processdesign - User contributions [en]

Desalination - Team E

2016-03-12T03:41:20Z

Bts511: /* Design Basis and Technical Approach */

Group E Corporation

Authors: Hassan Ali, Woo Soo Choe, Brett Sleyster, Jake Stolley

Instructors: Fengqi You, David Wegerer

March 11, 2016

=Executive Summary=

Outlined in this report is the project model report for the proposed desalination plant in Los Angeles, California which will provide drinking water to local citizens using reverse osmosis process. This location was chosen due to the region’s propensity for droughts in recent years. The fresh water resources in California have plummeted as water prices have risen 3-4% each year. This is only expected to continue. Other companies have worked deals with the Californian government to create desalination plants that provide a new fresh water drinking source. This suggests that there is a strong market here with potential growth.

The plant will have an output of 15 million gallons per day, which is about medium size relative to other plants. Thus the build time will be about one year. The plant will run for 30 years, with 350 days per year of use. Shutdown time will be used for cleaning and maintenance of equipment. The plant will be built on Long Beach, 500 meters from the coast nearby the Long Beach Water Reclamation Plant (LBWP). The input will withdraw 70.5 million gallons per day of seawater that will be used for reverse osmosis and dilution of brine.

Reverse osmosis was chosen as the desalination method for its low energy cost and flexibility. We designed a process that will pretreat the seawater to remove waste, colloids, bacteria, and membrane-damaging scalants. Once this seawater is treated is pumped to heavy pressures through a reverse osmosis system. After doing process optimization under various system parameter changes, we found a continuous, single stage, five element reverse osmosis process to be best. It produces a near 30% recovery. We reduce energy costs by including a pressure exchanger between our concentrate outlet and adjusted feed into the reverse osmosis system.

After performing mass and energy balances, we were able to size equipment and find our permeate composition. It turns out our process will meet L.A. standards for drinking water quality. However, post treatment will be done by the local LBWP to make the water taste better and be non-caustic to pipes. This will reduce “in-house” costs. Many environmental considerations were made, to meet our environmental standards. We meet EPA salinity requirements for our brine waste by diluting it. We also don’t use a lot of energy and thus reduce our fossil fuel consumption and CO2 generation. To be safe, we work well under max pressure limits for our equipment. We make sure we are ethical and meet all standards that our local county requires of us.

After an economic evaluation at a conservative well-water price of $700 per acre-foot, we realize that our revenue is way too low to cover our costs. Capital costs are especially large with regards to our large water tanks. We find that we reach a 10 year NPV breakpoint of $2104 per acre-foot after performing a sensitivity analysis. The Carlsbad Plant in San Diego worked out a deal for $2257 per acre-foot.

As we have met all design constraints for this project, we hope to convince policymakers to adopt our project at least $2104 per acre-foot water. If this deal is reached, we recommend that this project and design be funded. There is future work to be done to improve costs such as performing pump optimization. However, this work is minimal and won’t harm our bottom line if we reach our proposed price points.

__TOC__

=Introduction=

Due to a combination of various factors--including rising human population, climate change, increased energy consumption, and land erosion--the availability of fresh water as a resource for human consumption has been dwindling. This has had a negative impact on the world, but this impact has been especially felt in areas with natural shortages in fresh water. Desert regions, such as Saudi Arabia, have resorted to desalinating seawater at the coast and transporting up to their cities. This encompasses their entire freshwater system as very few natural water sources remain. In the United States there is a region which has shown increasing signs of following in the footsteps of Saudi Arabia and becoming fully reliant on desalination: Los Angeles, California.

=Location and Market Analysis=

In this project, Los Angeles was chosen as the location our plant, because the area has been subject to a severe drought for the last four years, as shown in Figure 1. The rapid decrease in the available water has caused permanent damages to the ground, decreased the amount of water in the reservoir from 45% to 25%, and reduced the amount of hydroelectric power produced in the state. Such deficiency in usable water has increased the water prices in Los Angeles by 6-7% annually for the last decade and is expected to rise by 3.4% annually for typical customers for the next five years. For customers who use more than 20000 gallons a year, the bill is expected to increase by 34% by 2021. This price increase is occurring because of both drought and maintenance. <ref>Pacific Institute: The California Drought. 2014.</ref>

[[File:Figure_1.jpg|right|200px|frame|Figure 1. California Drought Index.]]
As of March 2015 the Metropolitan Water District of Southern California was willing to pay as much as $700 per acre-foot of water to farmers in the Sacramento Valley, so we assume this will be the target price of the water produced by the this desalination plant.<ref>Hessel, Phil. “Dry Southern California Offers Northern Farmers Top Dollar for Water”. NBC News. March 18, 2015. </ref>

Currently, there is a plant in Carlsbad which supplies drinking water to the San Diego area. The plant can produce up to 50 million gallons per day and is the nation's largest water desalination plant. The water from this plant costs $2257 per acre-foot for the first 48000 acre-feet and about $2000 for any additional acre-foot. The county signed a contract with this plant to purchase the first 48000 acre-feet of water for the next 30 years.<ref>Carlsbad Desalination Plant. “Project Overview”. January 2016.</ref> Also, there is another desalination plant in Tampa that produces about 25 million gallons of water per day, and over 2000 other public and industrial plants with production capacity of over 300000 gallons per day. When conducting the economic analysis our plant, such numbers presented by other plants must be considered, so our project can yield positive profit.<ref>Leven, Rachel. “U.S. Desalination Industry Grows Since 2000; See as Essential to Meeting Supply Needs”. Bloomberg BNA. August 21, 2013.</ref>

The profitability of the project is naturally heavily dependent on the price of water in California. If the drought ends, water supply will increase and the water price will drop significantly, but if the drought continues, the price of water in California could rise even more rapidly and allow desalination plants to be more profitable.<ref>Southern California Public Radio. “Where is California water use decreasing?”. December 2015.</ref> It may be difficult to make this plant profitable without a guaranteed purchase or other subsidy from the local governments because the presence of pre-existing plants maintains the water price relatively low. However, external sources of profit such as guaranteed purchase or subsidy are expected because the ramifications of not having desalination plant or other sources of water may be significantly more expensive than the subsidizing a desalination plant assuming the drought continues.<ref>Weiser, Matt. “Could desalination solve California’s water problem?”. The Sacramento Bee. October 18, 2014.</ref>

=Design Basis and Technical Approach=

[[File:Figure_2.jpg|left|200px|frame|Figure 2. Vertical Well Feed]]
Considering the number of pre-existing desalination plant and the potential risk of the drought ending, our plant will be tentatively designed to produce 15 million gallons of drinkable water per day to homes and businesses in the Los Angeles county area. With a recovery rate of near 30%, and a requirement for diluting any brine, the input feed flowrate will be 70.5 million gallons per day. Following the similarly built Carlsbad plant, we expect the plant to last 30 years. The project building timeline is one year based on our relatively small plant size. We will run for 350 days a year while shutting down for cleaning and maintenance once every 3 months (based on DOW standards).<ref name = "Bates">Bates, Wayne. “Cleaning your RO”. Hydranautics. </ref> The plant will be built 500 meters from the coastline on Long Beach. A borehole will be drilled 10 meters into the ground. At this depth there should be seawater aquifers below the coast that regularly recharges itself with seawater from the ocean. A vertical well containing intake holes will be put into the borehole. The intake holes will be equipped with large size filters that can easily be replaced. At the surface, the well will be connected to a large-scale pump in a well pump house. From there, the seawater will be horizontally pumped into a seawater tank in the desalination plant. This will be considered the start of our process. The process ends after pretreatment, desalination, post treatment, and brine treatment are completed. Appendix 1 shows more details on the Design Basis.

The feed composition is assumed to be the standard seawater “average” based on the DOW Chemical ROSA 9.0 Technical Manual in Table 1. However, in reality the seawater composition at the plant site will have to be chemically tested. The final product will be based on L.A. standards for drinking water quality (Table 1). The specifications in the table are upper limits that we cannot exceed. We will aim to reach the salt level standards but any remineralization and pH treatment will be exported and paid for due to the added complexity and cost in making such a system in-house.

Due to the acidity of the desalinated water coming out of the reverse osmosis (RO) system (pH ~ 6), the pumps and pipes carrying the desalinated water for post treatment will likely corrode over time. We have decided to place the desalination plant by the Long Beach Water Reclamation Plant (about 50 meters away) to prevent pipe corrosion.

[[File:Table_1.jpg|center|300px|frame|Table 1. Concentration of the standard seawater, the output from the RO system, and LA standard.]]

=Desalination Overview and Alternatives=

We first began our research into the current landscape of the desalination industry, particularly of plants that were built in the United States. We discovered there were many processes but the main ones included: 
* Membrane Separation (Forward Osmosis and Reverse Osmosis)
* Thermal Separation (Multi Stage Flash Distillation)
* Phase Change Separation (Freeze Separation)
Each process was analyzed for their pros and cons. For our project design, we decided to go with reverse osmosis as our desalination process. This is because it is the most commonly used process in the industry, and for good reason. Unlike thermal separation, which requires huge amounts of energy to supply heat, reverse osmosis is very energy efficient. This reduces energy costs that would be spread across the plant’s lifetime.<ref name = "Republic">Republic of Mauritius NY. "Technical Aspects of Desalination Plant". Sustainable Sanitation and Water Management. </ref>

Forward osmosis is a newer technology being developed that also uses membranes to separate water and salt and may be more energy efficient overall than reverse osmosis. This is because forward osmosis doesn’t require a pump to pressurize seawater through the membranes, but rather leaves a higher salt concentrated solution (AKA draw solution) on the other side of the membrane for the seawater to be drawn into. Then a second step is required to regenerate the concentrated draw solution and produce purified water.<ref>Mallinson, Alissa. "Study Shows Forward Osmosis Desalination Not Energy Efficient." MIT News [Boston]. July 23, 2014.</ref> We decided to move away from forward osmosis because it is a process that is still being refined and developed for desalination and lacks the data, programs, and technologies that reverse osmosis users benefit from. Additionally, a recent study from MIT has claimed that due to the inherent nature of forward osmosis requiring that seawater be drawn to a higher concentrated salt solution before the water and salt is separated, it cannot possibly be more energy efficient than reverse osmosis.<ref>O'Hern, Sean C., Karnik, Rohit. Nanofiltration across Defect-Sealed Nanoporous Monolayer Graphene. American Chemical Society 15.5 (2014); 3254-260. </ref>

Other processes such as freeze separation were too restrictive in their feed and process demands and required a lot of capital. Ultimately, reverse osmosis is simply the best process that exists for desalination and has been ranked the best when performance is evaluated in terms of technical, environmental, and economic aspects. It is a process that is flexible in water quantity, operation start-up and shut-off, and site location.

Nevertheless there are disadvantages to reverse osmosis that must be considered. The feed seawater must not be too salty and too low quality. The pressure to pump the seawater is enormous. The construction time is long for large plants. The capital costs are also high when compared with other processes such as thermal separation. But all of these disadvantages are ultimately outweighed by the advantages reverse osmosis provides.<ref name = "Republic"/>

=Process Overview and Alternatives=

See Appendix 2 (Figure 1) for the Process Flow Diagram (PFD). The process itself is detailed below. The stream flow rates are in Appendix 2 (Table 1). We assumed 10 meters of pipe for every stream except the inlet and outlet streams of the entire PFD.

==Pretreatment==

After the 10 meter vertical well pumps seawater from the Long Beach coast, the feed is pumped 500 meters to a seawater tank that makes it easy to shut-down the system and reduce the large volume of flow. The level control keeps the seawater from overfilling the tank. From there, part of the seawater goes to brine treatment while the rest go to pretreatment. The purpose of pretreatment in the desalination process is to prevent equipment damage by removing particles and microorganisms from the water. A variety of pollutants -- including calcium carbonate, sulfates, silica, clay, bacteria, and other colloids and microorganisms -- can disrupt the reverse osmosis process. Media filtration uses fine grains such as sand or anthracite to filter the particles. We used a large filter to catch larger particle material (e.g. sand). However, simple filtration is not enough to prevent colloid fouling. An ultrafilter (i.e. microfiltration) is used to remove virtually all suspended matter, effectively preventing colloid fouling. The costs associated with ultrafiltration are not prohibitive in terms of the project’s budget.<ref name = "Desalination">Desalination Pre-treatment. Lenntech.</ref>

Reverse osmosis systems can also be disrupted by the growth of microbiological films of bacteria on the film and pipe walls. They are typically treated as colloids during pretreatment, and most organisms are removed from the water by ultrafiltration, but this does not remove the smallest of organic particles. DBNPA is the most effective biocide, but these aren’t often used and have to be removed from the water. Chlorination is an effective way to remove bacteria, but it can cause degradation of the membrane, and must be removed before the reverse osmosis occurs. In our process we will use UV irradiation to remove any organic material that passes through the microfilters.<ref name = "Desalination"/>

Scaling can reduce the flux across the membrane and can build up in pipe corners, causing backup. It is caused by inorganic salts -- specifically sulfates and calcium and magnesium carbonate -- present in the water. Scaling can be prevented and/or controlled with acids, antiscalants, or ion exchange softening. Antiscalants are good for removing sulfates, but not much else. Acidification will only be effective at removing CaCO3 from water but, the decreased pH may deteriorate the membrane and pipes, so it may require future research.<ref name = "Desalination"/> Ion exchange is very expensive and requires more capital to undergo, but is overall the more effective option.

The ion exchanger is a pretreatment system which acts as a method to fight membrane scaling in reverse osmosis. The ion exchanger works by using sodium carboxylate as a resin which will react with Chlorine and Magnesium ions, removing them from the feed. The sodium is substituted out and enters the feed stream, increasing the overall sodium level and TDS. The lack of magnesium and chlorine ions means that there is nothing to react with the carbonates to form scaling salts on the membranes.

==Desalination (Reverse Osmosis)==

Two pumps (P-105 & P-106) take the adjusted feed seawater up to 58 bar. However, before this pumping occurs, about 70% of the original stream gets separated and heads into a PX Pressure Exchanger. This separated stream later joins together with the pumped stream to enter the reverse osmosis system at the required pressure of 58 bar. The streams can combine together as they have equal pressure. It is the concentrate stream from the output of the reverse osmosis system that transfers its energy to the adjusted feed in the PX Pressure Exchanger.

[[File:Figure_3.jpg|left|300px|frame|Figure 3. Inside look of pressure exchanger]]
The pressure of the concentrate stream is about 57 bar and it is fed into an exchanger which works by using vessel that resembles a small pipe within a rotating wheel. The low pressure (LP) feed enters the vessel at 1 bar (see Figure 3). The feed turns and is pushed out into an outlet pipe by the high pressure (HP) concentrate stream at ~57 bar. The feed is now a HP stream via this exchanged energy. The now LP concentrate stream turns and is pushed out into an outlet pipe by the LP feed entering the vessel.

98% of the energy within the concentrate stream can be recovered via the PX Pressure Exchanger.<ref>Desalination Product and Applications Catalog 2013. “Advance Technologies for Seawater Reverse Osmosis and Brackish Water Applications”. Energy Recovery, Inc. 2012-2014.</ref >Thus, the now HP feed stream is not quite 58 bar and will need to be pumped by a small booster pump (P-107).

When designing a reverse osmosis system, we need to make decisions on what specific kind of reverse osmosis system will be implemented. When making the decision, we investigated the advantages of different types of reverse osmosis among the following options: batch vs. continuous process, number of stages, and number of elements.<ref>FILMTEC Reverse Osmosis Membranes. “Water & Process Solutions Technical Manual.” Dow Chemical Company.</ref> Our decision was made based on process optimization which is described later in this report. The end result of that optimization was to go with a single stage, continuous process with five elements in series.

=Post Treatment=

The permeate is sent to post treatment after the salt content is analyzed and controlled if too high or too low due to any unexpected errors. As stated above, the post treatment process will be done 50 meters from our plant by the Long Beach Water Reclamation Plant. The reason we aren’t performing post treatment “in-house” is due to the increased cost it would take for us to build and run the equipment for post treatment. We have the option to use the government’s water treatment facilities that already exist and our 15 million gallons per day output is a relatively small addition to the large Long Beach Water Reclamation Plant. We will assume any costs to post treatment will be pushed to the consumer in the price of purchasing the drinking water from the Long Beach Water Reclamation Plant.

Post treatment at any facility works to bring water to “taste standards”. The water that is the permeate of our process is already meeting L.A. drinking standards as shown above in Table 1 after our mass balance calculations around the reverse osmosis system (Appendix 3, Table 1). The post treatment increases the pH from 6.3 to 7.3 with the addition of base and then adds minerals like fluoride, calcium, and magnesium back in to improve taste.

==Brine Treatment and Byproduct Waste==

A natural consequence of a reverse osmosis procedure is that a lot of salt will be leftover from the membrane separation in a very concentrated form. This is called “brine” and will have to be taken care of. According to EPA requirements, the salt concentration of water entering the environment (via land or sea) cannot be higher than 40,000 ppm.8 The salt concentration of our brine is much higher (around 70,000 ppm). Thus, the brine must be removed or treated. We decided to treat the brine by diluting it as removing brine can be difficult and may require lots of land (e.g. evaporation ponds).

After the concentrate exchanges energy in the energy exchanger, it becomes a low pressure stream. It is pumped into a dilution tank and mixes with 18.2 million gallons per day of seawater from the inlet to become diluted brine which is pumped 600 meters out to sea.

The waste byproducts are the solid buildup on the filters and membranes. The filters and membranes also have a set lifetime and thus will also need to be removed. The waste is all non-hazardous and thus can be put in the municipal dump.

=Process Details and Equipment Sizing=

==Filters and UV Treatment [Pretreatment]==
The cost of our ultrafiltration system is estimated using data from the University of Tennessee at Knoxville.<ref>Qiang He and R. Bruce Robinson Membrane Filtration Systems. “Membrane Module Costs.” University of Tennessee at Knoxville. December 2014.</ref> It estimated the cost of the cost of the ultrafiltration $0.057 per 1000 gallons and the construction and capital costs are $0.39 per gallon per day. The cost of UV treatment is based off of units available from buyultraviolet.com for units that can treat about 450 gallons per minute.<ref>Atlantic Ultraviolet Corporation. “Megatron Water Disinfection System.”</ref> Due to the water having to be in relatively small pipes for treatment we do not expect the costs to scale very much, so we believe these units to be an accurate cost estimate.

==Ion Exchanger [Pretreatment] ==

The ion exchanger was sized using data from DOW. Our process will use 60 vessels with ion exchange beds of volume 5.73 m3. With a flow rate of 30 bed volumes/hour (10,312.2 m3/hr) the system of beds has a capacity of 125% of the feed. This allows for periodic regeneration of the resin, a weak acid cation resin (DOW MAC-3 Resin). This weak cation resin will remove cationic calcium and magnesium from the water to prevent scaling. The mass flow was calculated via a mass balance on the ion exchanger (Appendix 3).

==Reverse Osmosis System [Desalination]==

Each of the 2607 Reverse Osmosis SW30HRLE-440i [Seawater 30 High Rejection Low Energy 440 ft2 interlocking cap membrane] units have the same dimensions. A, the net length, is 1016 mm. B, the total length including the endcap, is 1029 mm. C, the diameter of the filter element, is 201mm. D, the diameter of the permeate pathway, is 29 mm. For a single stage, we have five of the elements in series so the overall length of a single stage is expected to be roughly 5145 mm. The membrane material is polyamide thin film composite. It will last three years.
Operating Conditions [ROSA 9.0]: 
* Temperature: 30oC 
* Pressure: 58 bar 
* Flow Rate: 52.3 Million gal per day 
* Recovery: 28.7%

[[File:Figure_4.png|center|400px|frame|Figure 4. Diagram of Dow FILMTEC SW30HRLE-440i]]
Cleaning Schedule: 
* Conservative Estimate: Once every three months (three days available for cleaning)<ref name = "Bates"/> 
* DOW FILMTEC Cleaning Conditions 
** The normalized permeate flow drops 10%
* The normalized salt passage increases 5 - 10%

==Seawater Holding Tank and Dilution Tank [Brine Treatment]==

The initial seawater holding tank works as a buffer to prevent surges and help perform shutdowns quickly and easily. We expect to run continuously for 350 days a year with the rest of the available time spread out and used for shutting down the plant for cleaning membranes, filters, and the fixing of unexpected errors. There is a big tradeoff between having a holding tank with high retention time and having a holding tank with low cost. With that in mind, we erred on the side of lowering capital cost and went with a retention time of one hour. The volume could be found by multiplying the retention time with the feed flow rate (Stream 1). We assumed that a good height to width ratio for a large water tank would be 4:1.

A mass balance was done around the dilution tank to find what amount of seawater would be necessary to dilute the brine to 45,000 ppm salinity (Appendix 3; Table 3). With this the feed flow rate into the tank could be found. It was found that it takes about 20 minutes of retention time for good dilution of the brine in seawater.8 We again assumed that a good height to width ratio for a large water tank would be 4:1.

We assumed a corrosion allowance of 4 mm wall thickness. We calculated wall thickness using vessel pressure, hoop stress, longitudinal stress, and weld factor. Each tank is made with SS316 to protect against corrosion.<ref>Towler G, Sinnott R. Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013.</ref>

<center>
'''Table 2.''' Tank sizes.
[[File:Table_2.jpg|center|600px]]
</center>

=Process Optimization=
<center> '''Table 3.'''Optimization Criteria
[[File:Table_3.jpg|center|400px]]
</center>
In order to optimize the reverse osmosis system, numerous simulations were run. Three representative cases are reported as shown above (Table 3). The objective of the optimization was to minimize the required water price to break even in 10 years. In order to achieve this goal, the balance between the capital cost from the number of the filter elements and utility cost from energy consumption was weighed. Before the energy recovery system was implemented, the general trend was that if we have more filter elements in series, overall recovery increases but recovery rate per filter decreases, and we need more filter elements. While increasing the elements in series increase the required number of filters, and thereby increases the capital cost, we require less energy to produce the same amount of water. After the energy recovery system is implemented, the correlation between the number of elements in series and the required energy becomes more complicated because high pressure concentrate is used to pressurize the feed. However, the trade-off between the utility cost and capital cost still exists. Of the three cases shown above, single stage, five elements RO system turned out to be the most optimal solution because single stage, single element system required higher utility cost and two stage system slightly did not meet the LA standard. Detailed information which was used for the optimization is available in Appendix 3 and in Appendix 6.

=Ethical, Environmental, and Societal Considerations=

There are a few key considerations we have followed to meet ethical, environmental, and societal guidelines and requirements: 
* Meet L.A. Drinking Water Standards
** Correct ion concentration
* Moved near water treatment plant
** Pipes will not corrode from low pH
* Working below pressure limits
* Water will be locally served and meet taste standards
** Via Post Treatment mineralization
* Follows EPA Environmental Salinity Requirements
* Reduced Energy Requirements
** PX Pressure Exchanger
** Optimized RO System
** Reverse Osmosis is itself Energy Saving

For further environmental considerations, see effluent summary in Appendix 4.

=Economic Analysis=

The pricing of the majority of the process was completed using the Aspen Economic Analyzer, however we had determine costs for the ultrafiltration, reverse osmosis membranes, and the ultraviolet treatment was completed separately. This led to a total ISBL of $56 MM, OSBL of $22.4 MM, Engineering Costs of $22.4 MM and a Contingency of $7.8 MM. We also chose a working Capital of $10 MM. This working capital is sufficient because we have no raw materials costs. We also had FCOP of $5.2 MM and VCOP of $3.5 MM.

We originally were planning on pricing our water at $700 per acre-foot, but after doing economic analysis we had a 10 year NPV of $-76.2 MM. After determining did a sensitivity analysis and determined that if our water was priced at $2104 per acre-foot we could have an NPV of $0 after 10 years and if we were able to get the same price of $2257 per acre-foot as the plant in Carlsbad we would have a 10 year NPV of $8.3 MM which increases to $41.1 MM at 20 years. For more detailed data about the economic analysis see Appendix 6.

=Conclusion and Recommendations=

Our mission was to create an economically viable process that will desalinate water for human consumption in the L.A. County. We have found after research and optimization that the best process to undergo desalination is a continuous single stage, five elements reverse osmosis system. Our process can meet the required L.A. drinking water standard and we have set up plans for post treatment at a local water reclamation site. While meeting the drinking water standards, we also met our ethical, environmental, and social standards.

Nevertheless, at a sale price of $700 per acre-foot, our proposed desalination plant in Los Angeles, California looks to be a very expensive project to undergo. Without any sort of government assistance to carry the heavy capital costs of such a project, we would suggest the plant not be built. However, after performing a sensitivity analysis at higher (and less conservative) price points we reached a breakeven point in 10 years with a sale price of $2104 per acre-foot of drinking water. Research has shown that the droughts striking California look only to get worse in the near future. With limited freshwater resources, there is no option other than to consider desalination. Therefore it is reasonable to expect we could get near the same price deal as local desalination plants like the Carlsbad Plant ($2257 per acre-foot) As we have met all design constraints for this project, we hope to convince policymakers to adopt our project at least $2104 per acre-foot water. If this deal is reached, we recommend that this project and design be funded.

However, there are ways to improve our design with future work. Right now all the pipes diameters are based on having a flowrate of 1.5 m/s. But in reality, we’d need to use standard piping sizes to get more accurate pricing and energy costs. Also, in reality we would be purchasing some of these vessels in bulk quantities (i.e. bulk filters) so the price could be lowered with negotiation and economies of scale. We also have pumps that are providing very little pressure for their high cost. We can remove these pumps and simply run at higher pressures earlier. We can potentially reduce early energy costs by building a “floating” desalination plant in the sea to reduce pumping distance from the well.

=Appendices=

==Appendix 1: Design Basis==
[[File:Appendix_1.png|center|Appendix_1.png]]

==Appendix 2: Process Flow Diagram==
[[File:Appendix_2.png|center]]
==Appendix 3: Mass and Energy Balance Tables==
'''Table 1'''. Mass balance on 1 stage 5 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T1.png|center|500px|]]

'''Table 2'''. Mass balance on 1 stage 1 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T2.png|center|500px|]]

'''Table 3'''. Mass balance on 2 stage 5,5 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T3.png|center|600px|]]

'''Table 4'''. 1 stage 5 elements CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T4.png|center|300px|]]

'''Table 5'''. 1 stage 1 element CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T5.png|center|300px|]]

'''Table 6'''. 2 stages 5,5 elements CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T6.png|center|300px|]]

'''Table 7'''. Energy balance on 1 stage 5 elements RO system
[[File:Appendix_3_T7.png|center|800px|]]

==Appendix 4: Effluent Summary==
[[File:Appendix_4.png|center]]

==Appendix 5: Cost of Production==
[[File:Appendix_5.png|center]]

==Appendix 6: Economic Anlaysis==
===Case I ===
Water is $700; RO with 1 stage and 5 elements
[[File:Appendix_6_C1.png|center]]

===Case II===
Water is $1400; RO with 1 stage and 5 elements
[[File:Appendix_6_C2.png|center]]

===Case III: Carlsbad ===
Water is $2257; RO with 1 stage and 5 elements
[[File:Appendix_6_C3.png|center]]

===Case IV: Breakeven (NPV = 0 at 10 years)===
Water is $2104; RO with 1 stage and 5 elements
[[File:Appendix_6_C4.png|center]]

===Case V: Breakeven (NPV = 0 at 10 years)===
Water is $2112; RO with 2 stages and 5,5 elements
[[File:Appendix_6_C5.png|center]]

===Case VI: Breakeven (NPV = 0 at 10 years)===
Water is $2138 with 1 stage and 1 element
[[File:Appendix_6_C6.png|center]]

==Appendix 7: ROSA Simulation Outputs==

'''Table 1'''. System detail of 1 stage 5 elements RO system
[[File:Appendix_7_T1.png|center|500px|]]

'''Table 2'''. System detail of 1 stage 1 elements RO system
[[File:Appendix_7_T2.png|center|500px|]]

'''Table 3'''. System detail of 2 stage 5,5 elements system
[[File:Appendix_7_T3.png|center|500px|]]

'''Table 4'''. 1 stage 5 elements RO system concentrations and pH in each stream
[[File:Appendix_7_T4.png|center|500px|]]

'''Table 5'''. 1 stage 1 element RO system concentration and pH in each stream
[[File:Appendix_7_T5.png|center|500px|]]

'''Table 6'''. 2 stage 5,5 elements RO system concentrations and pH in each stream
[[File:Appendix_7_T6.png|center|700px|]]

=References=
<references/>

Desalination - Team E

2016-03-12T03:40:52Z

Bts511: /* Design Basis and Technical Approach */

Group E Corporation

Authors: Hassan Ali, Woo Soo Choe, Brett Sleyster, Jake Stolley

Instructors: Fengqi You, David Wegerer

March 11, 2016

=Executive Summary=

Outlined in this report is the project model report for the proposed desalination plant in Los Angeles, California which will provide drinking water to local citizens using reverse osmosis process. This location was chosen due to the region’s propensity for droughts in recent years. The fresh water resources in California have plummeted as water prices have risen 3-4% each year. This is only expected to continue. Other companies have worked deals with the Californian government to create desalination plants that provide a new fresh water drinking source. This suggests that there is a strong market here with potential growth.

The plant will have an output of 15 million gallons per day, which is about medium size relative to other plants. Thus the build time will be about one year. The plant will run for 30 years, with 350 days per year of use. Shutdown time will be used for cleaning and maintenance of equipment. The plant will be built on Long Beach, 500 meters from the coast nearby the Long Beach Water Reclamation Plant (LBWP). The input will withdraw 70.5 million gallons per day of seawater that will be used for reverse osmosis and dilution of brine.

Reverse osmosis was chosen as the desalination method for its low energy cost and flexibility. We designed a process that will pretreat the seawater to remove waste, colloids, bacteria, and membrane-damaging scalants. Once this seawater is treated is pumped to heavy pressures through a reverse osmosis system. After doing process optimization under various system parameter changes, we found a continuous, single stage, five element reverse osmosis process to be best. It produces a near 30% recovery. We reduce energy costs by including a pressure exchanger between our concentrate outlet and adjusted feed into the reverse osmosis system.

After performing mass and energy balances, we were able to size equipment and find our permeate composition. It turns out our process will meet L.A. standards for drinking water quality. However, post treatment will be done by the local LBWP to make the water taste better and be non-caustic to pipes. This will reduce “in-house” costs. Many environmental considerations were made, to meet our environmental standards. We meet EPA salinity requirements for our brine waste by diluting it. We also don’t use a lot of energy and thus reduce our fossil fuel consumption and CO2 generation. To be safe, we work well under max pressure limits for our equipment. We make sure we are ethical and meet all standards that our local county requires of us.

After an economic evaluation at a conservative well-water price of $700 per acre-foot, we realize that our revenue is way too low to cover our costs. Capital costs are especially large with regards to our large water tanks. We find that we reach a 10 year NPV breakpoint of $2104 per acre-foot after performing a sensitivity analysis. The Carlsbad Plant in San Diego worked out a deal for $2257 per acre-foot.

As we have met all design constraints for this project, we hope to convince policymakers to adopt our project at least $2104 per acre-foot water. If this deal is reached, we recommend that this project and design be funded. There is future work to be done to improve costs such as performing pump optimization. However, this work is minimal and won’t harm our bottom line if we reach our proposed price points.

__TOC__

=Introduction=

Due to a combination of various factors--including rising human population, climate change, increased energy consumption, and land erosion--the availability of fresh water as a resource for human consumption has been dwindling. This has had a negative impact on the world, but this impact has been especially felt in areas with natural shortages in fresh water. Desert regions, such as Saudi Arabia, have resorted to desalinating seawater at the coast and transporting up to their cities. This encompasses their entire freshwater system as very few natural water sources remain. In the United States there is a region which has shown increasing signs of following in the footsteps of Saudi Arabia and becoming fully reliant on desalination: Los Angeles, California.

=Location and Market Analysis=

In this project, Los Angeles was chosen as the location our plant, because the area has been subject to a severe drought for the last four years, as shown in Figure 1. The rapid decrease in the available water has caused permanent damages to the ground, decreased the amount of water in the reservoir from 45% to 25%, and reduced the amount of hydroelectric power produced in the state. Such deficiency in usable water has increased the water prices in Los Angeles by 6-7% annually for the last decade and is expected to rise by 3.4% annually for typical customers for the next five years. For customers who use more than 20000 gallons a year, the bill is expected to increase by 34% by 2021. This price increase is occurring because of both drought and maintenance. <ref>Pacific Institute: The California Drought. 2014.</ref>

[[File:Figure_1.jpg|right|200px|frame|Figure 1. California Drought Index.]]
As of March 2015 the Metropolitan Water District of Southern California was willing to pay as much as $700 per acre-foot of water to farmers in the Sacramento Valley, so we assume this will be the target price of the water produced by the this desalination plant.<ref>Hessel, Phil. “Dry Southern California Offers Northern Farmers Top Dollar for Water”. NBC News. March 18, 2015. </ref>

Currently, there is a plant in Carlsbad which supplies drinking water to the San Diego area. The plant can produce up to 50 million gallons per day and is the nation's largest water desalination plant. The water from this plant costs $2257 per acre-foot for the first 48000 acre-feet and about $2000 for any additional acre-foot. The county signed a contract with this plant to purchase the first 48000 acre-feet of water for the next 30 years.<ref>Carlsbad Desalination Plant. “Project Overview”. January 2016.</ref> Also, there is another desalination plant in Tampa that produces about 25 million gallons of water per day, and over 2000 other public and industrial plants with production capacity of over 300000 gallons per day. When conducting the economic analysis our plant, such numbers presented by other plants must be considered, so our project can yield positive profit.<ref>Leven, Rachel. “U.S. Desalination Industry Grows Since 2000; See as Essential to Meeting Supply Needs”. Bloomberg BNA. August 21, 2013.</ref>

The profitability of the project is naturally heavily dependent on the price of water in California. If the drought ends, water supply will increase and the water price will drop significantly, but if the drought continues, the price of water in California could rise even more rapidly and allow desalination plants to be more profitable.<ref>Southern California Public Radio. “Where is California water use decreasing?”. December 2015.</ref> It may be difficult to make this plant profitable without a guaranteed purchase or other subsidy from the local governments because the presence of pre-existing plants maintains the water price relatively low. However, external sources of profit such as guaranteed purchase or subsidy are expected because the ramifications of not having desalination plant or other sources of water may be significantly more expensive than the subsidizing a desalination plant assuming the drought continues.<ref>Weiser, Matt. “Could desalination solve California’s water problem?”. The Sacramento Bee. October 18, 2014.</ref>

=Design Basis and Technical Approach=

[[File:Figure_2.jpg|left|200px|frame|Figure 2. Vertical Well Feed]]
Considering the number of pre-existing desalination plant and the potential risk of the drought ending, our plant will be tentatively designed to produce 15 million gallons of drinkable water per day to homes and businesses in the Los Angeles county area. With a recovery rate of near 30%, and a requirement for diluting any brine, the input feed flowrate will be 70.5 million gallons per day. Following the similarly built Carlsbad plant, we expect the plant to last 30 years. The project building timeline is one year based on our relatively small plant size. We will run for 350 days a year while shutting down for cleaning and maintenance once every 3 months (based on DOW standards).<ref name = "Bates">Bates, Wayne. “Cleaning your RO”. Hydranautics. </ref> The plant will be built 500 meters from the coastline on Long Beach. A borehole will be drilled 10 meters into the ground. At this depth there should be seawater aquifers below the coast that regularly recharges itself with seawater from the ocean. A vertical well containing intake holes will be put into the borehole. The intake holes will be equipped with large size filters that can easily be replaced. At the surface, the well will be connected to a large-scale pump in a well pump house. From there, the seawater will be horizontally pumped into a seawater tank in the desalination plant. This will be considered the start of our process. The process ends after pretreatment, desalination, post treatment, and brine treatment are completed. Appendix 1 shows more details on the Design Basis.

The feed composition is assumed to be the standard seawater “average” based on the DOW Chemical ROSA 9.0 Technical Manual in Table 1. However, in reality the seawater composition at the plant site will have to be chemically tested. The final product will be based on L.A. standards for drinking water quality (Table 1). The specifications in the table are upper limits that we cannot exceed. We will aim to reach the salt level standards but any remineralization and pH treatment will be exported and paid for due to the added complexity and cost in making such a system in-house.

Due to the acidity of the desalinated water coming out of the reverse osmosis (RO) system (pH ~ 6), the pumps and pipes carrying the desalinated water for post treatment will likely corrode over time. We have decided to place the desalination plant by the Long Beach Water Reclamation Plant (about 50 meters away) to prevent pipe corrosion.

[[File:Table_1.jpg|center|800px|thumb|Table 1. Concentration of the standard seawater, the output from the RO system, and LA standard.]]

=Desalination Overview and Alternatives=

We first began our research into the current landscape of the desalination industry, particularly of plants that were built in the United States. We discovered there were many processes but the main ones included: 
* Membrane Separation (Forward Osmosis and Reverse Osmosis)
* Thermal Separation (Multi Stage Flash Distillation)
* Phase Change Separation (Freeze Separation)
Each process was analyzed for their pros and cons. For our project design, we decided to go with reverse osmosis as our desalination process. This is because it is the most commonly used process in the industry, and for good reason. Unlike thermal separation, which requires huge amounts of energy to supply heat, reverse osmosis is very energy efficient. This reduces energy costs that would be spread across the plant’s lifetime.<ref name = "Republic">Republic of Mauritius NY. "Technical Aspects of Desalination Plant". Sustainable Sanitation and Water Management. </ref>

Forward osmosis is a newer technology being developed that also uses membranes to separate water and salt and may be more energy efficient overall than reverse osmosis. This is because forward osmosis doesn’t require a pump to pressurize seawater through the membranes, but rather leaves a higher salt concentrated solution (AKA draw solution) on the other side of the membrane for the seawater to be drawn into. Then a second step is required to regenerate the concentrated draw solution and produce purified water.<ref>Mallinson, Alissa. "Study Shows Forward Osmosis Desalination Not Energy Efficient." MIT News [Boston]. July 23, 2014.</ref> We decided to move away from forward osmosis because it is a process that is still being refined and developed for desalination and lacks the data, programs, and technologies that reverse osmosis users benefit from. Additionally, a recent study from MIT has claimed that due to the inherent nature of forward osmosis requiring that seawater be drawn to a higher concentrated salt solution before the water and salt is separated, it cannot possibly be more energy efficient than reverse osmosis.<ref>O'Hern, Sean C., Karnik, Rohit. Nanofiltration across Defect-Sealed Nanoporous Monolayer Graphene. American Chemical Society 15.5 (2014); 3254-260. </ref>

Other processes such as freeze separation were too restrictive in their feed and process demands and required a lot of capital. Ultimately, reverse osmosis is simply the best process that exists for desalination and has been ranked the best when performance is evaluated in terms of technical, environmental, and economic aspects. It is a process that is flexible in water quantity, operation start-up and shut-off, and site location.

Nevertheless there are disadvantages to reverse osmosis that must be considered. The feed seawater must not be too salty and too low quality. The pressure to pump the seawater is enormous. The construction time is long for large plants. The capital costs are also high when compared with other processes such as thermal separation. But all of these disadvantages are ultimately outweighed by the advantages reverse osmosis provides.<ref name = "Republic"/>

=Process Overview and Alternatives=

See Appendix 2 (Figure 1) for the Process Flow Diagram (PFD). The process itself is detailed below. The stream flow rates are in Appendix 2 (Table 1). We assumed 10 meters of pipe for every stream except the inlet and outlet streams of the entire PFD.

==Pretreatment==

After the 10 meter vertical well pumps seawater from the Long Beach coast, the feed is pumped 500 meters to a seawater tank that makes it easy to shut-down the system and reduce the large volume of flow. The level control keeps the seawater from overfilling the tank. From there, part of the seawater goes to brine treatment while the rest go to pretreatment. The purpose of pretreatment in the desalination process is to prevent equipment damage by removing particles and microorganisms from the water. A variety of pollutants -- including calcium carbonate, sulfates, silica, clay, bacteria, and other colloids and microorganisms -- can disrupt the reverse osmosis process. Media filtration uses fine grains such as sand or anthracite to filter the particles. We used a large filter to catch larger particle material (e.g. sand). However, simple filtration is not enough to prevent colloid fouling. An ultrafilter (i.e. microfiltration) is used to remove virtually all suspended matter, effectively preventing colloid fouling. The costs associated with ultrafiltration are not prohibitive in terms of the project’s budget.<ref name = "Desalination">Desalination Pre-treatment. Lenntech.</ref>

Reverse osmosis systems can also be disrupted by the growth of microbiological films of bacteria on the film and pipe walls. They are typically treated as colloids during pretreatment, and most organisms are removed from the water by ultrafiltration, but this does not remove the smallest of organic particles. DBNPA is the most effective biocide, but these aren’t often used and have to be removed from the water. Chlorination is an effective way to remove bacteria, but it can cause degradation of the membrane, and must be removed before the reverse osmosis occurs. In our process we will use UV irradiation to remove any organic material that passes through the microfilters.<ref name = "Desalination"/>

Scaling can reduce the flux across the membrane and can build up in pipe corners, causing backup. It is caused by inorganic salts -- specifically sulfates and calcium and magnesium carbonate -- present in the water. Scaling can be prevented and/or controlled with acids, antiscalants, or ion exchange softening. Antiscalants are good for removing sulfates, but not much else. Acidification will only be effective at removing CaCO3 from water but, the decreased pH may deteriorate the membrane and pipes, so it may require future research.<ref name = "Desalination"/> Ion exchange is very expensive and requires more capital to undergo, but is overall the more effective option.

The ion exchanger is a pretreatment system which acts as a method to fight membrane scaling in reverse osmosis. The ion exchanger works by using sodium carboxylate as a resin which will react with Chlorine and Magnesium ions, removing them from the feed. The sodium is substituted out and enters the feed stream, increasing the overall sodium level and TDS. The lack of magnesium and chlorine ions means that there is nothing to react with the carbonates to form scaling salts on the membranes.

==Desalination (Reverse Osmosis)==

Two pumps (P-105 & P-106) take the adjusted feed seawater up to 58 bar. However, before this pumping occurs, about 70% of the original stream gets separated and heads into a PX Pressure Exchanger. This separated stream later joins together with the pumped stream to enter the reverse osmosis system at the required pressure of 58 bar. The streams can combine together as they have equal pressure. It is the concentrate stream from the output of the reverse osmosis system that transfers its energy to the adjusted feed in the PX Pressure Exchanger.

[[File:Figure_3.jpg|left|300px|frame|Figure 3. Inside look of pressure exchanger]]
The pressure of the concentrate stream is about 57 bar and it is fed into an exchanger which works by using vessel that resembles a small pipe within a rotating wheel. The low pressure (LP) feed enters the vessel at 1 bar (see Figure 3). The feed turns and is pushed out into an outlet pipe by the high pressure (HP) concentrate stream at ~57 bar. The feed is now a HP stream via this exchanged energy. The now LP concentrate stream turns and is pushed out into an outlet pipe by the LP feed entering the vessel.

98% of the energy within the concentrate stream can be recovered via the PX Pressure Exchanger.<ref>Desalination Product and Applications Catalog 2013. “Advance Technologies for Seawater Reverse Osmosis and Brackish Water Applications”. Energy Recovery, Inc. 2012-2014.</ref >Thus, the now HP feed stream is not quite 58 bar and will need to be pumped by a small booster pump (P-107).

When designing a reverse osmosis system, we need to make decisions on what specific kind of reverse osmosis system will be implemented. When making the decision, we investigated the advantages of different types of reverse osmosis among the following options: batch vs. continuous process, number of stages, and number of elements.<ref>FILMTEC Reverse Osmosis Membranes. “Water & Process Solutions Technical Manual.” Dow Chemical Company.</ref> Our decision was made based on process optimization which is described later in this report. The end result of that optimization was to go with a single stage, continuous process with five elements in series.

=Post Treatment=

The permeate is sent to post treatment after the salt content is analyzed and controlled if too high or too low due to any unexpected errors. As stated above, the post treatment process will be done 50 meters from our plant by the Long Beach Water Reclamation Plant. The reason we aren’t performing post treatment “in-house” is due to the increased cost it would take for us to build and run the equipment for post treatment. We have the option to use the government’s water treatment facilities that already exist and our 15 million gallons per day output is a relatively small addition to the large Long Beach Water Reclamation Plant. We will assume any costs to post treatment will be pushed to the consumer in the price of purchasing the drinking water from the Long Beach Water Reclamation Plant.

Post treatment at any facility works to bring water to “taste standards”. The water that is the permeate of our process is already meeting L.A. drinking standards as shown above in Table 1 after our mass balance calculations around the reverse osmosis system (Appendix 3, Table 1). The post treatment increases the pH from 6.3 to 7.3 with the addition of base and then adds minerals like fluoride, calcium, and magnesium back in to improve taste.

==Brine Treatment and Byproduct Waste==

A natural consequence of a reverse osmosis procedure is that a lot of salt will be leftover from the membrane separation in a very concentrated form. This is called “brine” and will have to be taken care of. According to EPA requirements, the salt concentration of water entering the environment (via land or sea) cannot be higher than 40,000 ppm.8 The salt concentration of our brine is much higher (around 70,000 ppm). Thus, the brine must be removed or treated. We decided to treat the brine by diluting it as removing brine can be difficult and may require lots of land (e.g. evaporation ponds).

After the concentrate exchanges energy in the energy exchanger, it becomes a low pressure stream. It is pumped into a dilution tank and mixes with 18.2 million gallons per day of seawater from the inlet to become diluted brine which is pumped 600 meters out to sea.

The waste byproducts are the solid buildup on the filters and membranes. The filters and membranes also have a set lifetime and thus will also need to be removed. The waste is all non-hazardous and thus can be put in the municipal dump.

=Process Details and Equipment Sizing=

==Filters and UV Treatment [Pretreatment]==
The cost of our ultrafiltration system is estimated using data from the University of Tennessee at Knoxville.<ref>Qiang He and R. Bruce Robinson Membrane Filtration Systems. “Membrane Module Costs.” University of Tennessee at Knoxville. December 2014.</ref> It estimated the cost of the cost of the ultrafiltration $0.057 per 1000 gallons and the construction and capital costs are $0.39 per gallon per day. The cost of UV treatment is based off of units available from buyultraviolet.com for units that can treat about 450 gallons per minute.<ref>Atlantic Ultraviolet Corporation. “Megatron Water Disinfection System.”</ref> Due to the water having to be in relatively small pipes for treatment we do not expect the costs to scale very much, so we believe these units to be an accurate cost estimate.

==Ion Exchanger [Pretreatment] ==

The ion exchanger was sized using data from DOW. Our process will use 60 vessels with ion exchange beds of volume 5.73 m3. With a flow rate of 30 bed volumes/hour (10,312.2 m3/hr) the system of beds has a capacity of 125% of the feed. This allows for periodic regeneration of the resin, a weak acid cation resin (DOW MAC-3 Resin). This weak cation resin will remove cationic calcium and magnesium from the water to prevent scaling. The mass flow was calculated via a mass balance on the ion exchanger (Appendix 3).

==Reverse Osmosis System [Desalination]==

Each of the 2607 Reverse Osmosis SW30HRLE-440i [Seawater 30 High Rejection Low Energy 440 ft2 interlocking cap membrane] units have the same dimensions. A, the net length, is 1016 mm. B, the total length including the endcap, is 1029 mm. C, the diameter of the filter element, is 201mm. D, the diameter of the permeate pathway, is 29 mm. For a single stage, we have five of the elements in series so the overall length of a single stage is expected to be roughly 5145 mm. The membrane material is polyamide thin film composite. It will last three years.
Operating Conditions [ROSA 9.0]: 
* Temperature: 30oC 
* Pressure: 58 bar 
* Flow Rate: 52.3 Million gal per day 
* Recovery: 28.7%

[[File:Figure_4.png|center|400px|frame|Figure 4. Diagram of Dow FILMTEC SW30HRLE-440i]]
Cleaning Schedule: 
* Conservative Estimate: Once every three months (three days available for cleaning)<ref name = "Bates"/> 
* DOW FILMTEC Cleaning Conditions 
** The normalized permeate flow drops 10%
* The normalized salt passage increases 5 - 10%

==Seawater Holding Tank and Dilution Tank [Brine Treatment]==

The initial seawater holding tank works as a buffer to prevent surges and help perform shutdowns quickly and easily. We expect to run continuously for 350 days a year with the rest of the available time spread out and used for shutting down the plant for cleaning membranes, filters, and the fixing of unexpected errors. There is a big tradeoff between having a holding tank with high retention time and having a holding tank with low cost. With that in mind, we erred on the side of lowering capital cost and went with a retention time of one hour. The volume could be found by multiplying the retention time with the feed flow rate (Stream 1). We assumed that a good height to width ratio for a large water tank would be 4:1.

A mass balance was done around the dilution tank to find what amount of seawater would be necessary to dilute the brine to 45,000 ppm salinity (Appendix 3; Table 3). With this the feed flow rate into the tank could be found. It was found that it takes about 20 minutes of retention time for good dilution of the brine in seawater.8 We again assumed that a good height to width ratio for a large water tank would be 4:1.

We assumed a corrosion allowance of 4 mm wall thickness. We calculated wall thickness using vessel pressure, hoop stress, longitudinal stress, and weld factor. Each tank is made with SS316 to protect against corrosion.<ref>Towler G, Sinnott R. Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013.</ref>

<center>
'''Table 2.''' Tank sizes.
[[File:Table_2.jpg|center|600px]]
</center>

=Process Optimization=
<center> '''Table 3.'''Optimization Criteria
[[File:Table_3.jpg|center|400px]]
</center>
In order to optimize the reverse osmosis system, numerous simulations were run. Three representative cases are reported as shown above (Table 3). The objective of the optimization was to minimize the required water price to break even in 10 years. In order to achieve this goal, the balance between the capital cost from the number of the filter elements and utility cost from energy consumption was weighed. Before the energy recovery system was implemented, the general trend was that if we have more filter elements in series, overall recovery increases but recovery rate per filter decreases, and we need more filter elements. While increasing the elements in series increase the required number of filters, and thereby increases the capital cost, we require less energy to produce the same amount of water. After the energy recovery system is implemented, the correlation between the number of elements in series and the required energy becomes more complicated because high pressure concentrate is used to pressurize the feed. However, the trade-off between the utility cost and capital cost still exists. Of the three cases shown above, single stage, five elements RO system turned out to be the most optimal solution because single stage, single element system required higher utility cost and two stage system slightly did not meet the LA standard. Detailed information which was used for the optimization is available in Appendix 3 and in Appendix 6.

=Ethical, Environmental, and Societal Considerations=

There are a few key considerations we have followed to meet ethical, environmental, and societal guidelines and requirements: 
* Meet L.A. Drinking Water Standards
** Correct ion concentration
* Moved near water treatment plant
** Pipes will not corrode from low pH
* Working below pressure limits
* Water will be locally served and meet taste standards
** Via Post Treatment mineralization
* Follows EPA Environmental Salinity Requirements
* Reduced Energy Requirements
** PX Pressure Exchanger
** Optimized RO System
** Reverse Osmosis is itself Energy Saving

For further environmental considerations, see effluent summary in Appendix 4.

=Economic Analysis=

The pricing of the majority of the process was completed using the Aspen Economic Analyzer, however we had determine costs for the ultrafiltration, reverse osmosis membranes, and the ultraviolet treatment was completed separately. This led to a total ISBL of $56 MM, OSBL of $22.4 MM, Engineering Costs of $22.4 MM and a Contingency of $7.8 MM. We also chose a working Capital of $10 MM. This working capital is sufficient because we have no raw materials costs. We also had FCOP of $5.2 MM and VCOP of $3.5 MM.

We originally were planning on pricing our water at $700 per acre-foot, but after doing economic analysis we had a 10 year NPV of $-76.2 MM. After determining did a sensitivity analysis and determined that if our water was priced at $2104 per acre-foot we could have an NPV of $0 after 10 years and if we were able to get the same price of $2257 per acre-foot as the plant in Carlsbad we would have a 10 year NPV of $8.3 MM which increases to $41.1 MM at 20 years. For more detailed data about the economic analysis see Appendix 6.

=Conclusion and Recommendations=

Our mission was to create an economically viable process that will desalinate water for human consumption in the L.A. County. We have found after research and optimization that the best process to undergo desalination is a continuous single stage, five elements reverse osmosis system. Our process can meet the required L.A. drinking water standard and we have set up plans for post treatment at a local water reclamation site. While meeting the drinking water standards, we also met our ethical, environmental, and social standards.

Nevertheless, at a sale price of $700 per acre-foot, our proposed desalination plant in Los Angeles, California looks to be a very expensive project to undergo. Without any sort of government assistance to carry the heavy capital costs of such a project, we would suggest the plant not be built. However, after performing a sensitivity analysis at higher (and less conservative) price points we reached a breakeven point in 10 years with a sale price of $2104 per acre-foot of drinking water. Research has shown that the droughts striking California look only to get worse in the near future. With limited freshwater resources, there is no option other than to consider desalination. Therefore it is reasonable to expect we could get near the same price deal as local desalination plants like the Carlsbad Plant ($2257 per acre-foot) As we have met all design constraints for this project, we hope to convince policymakers to adopt our project at least $2104 per acre-foot water. If this deal is reached, we recommend that this project and design be funded.

However, there are ways to improve our design with future work. Right now all the pipes diameters are based on having a flowrate of 1.5 m/s. But in reality, we’d need to use standard piping sizes to get more accurate pricing and energy costs. Also, in reality we would be purchasing some of these vessels in bulk quantities (i.e. bulk filters) so the price could be lowered with negotiation and economies of scale. We also have pumps that are providing very little pressure for their high cost. We can remove these pumps and simply run at higher pressures earlier. We can potentially reduce early energy costs by building a “floating” desalination plant in the sea to reduce pumping distance from the well.

=Appendices=

==Appendix 1: Design Basis==
[[File:Appendix_1.png|center|Appendix_1.png]]

==Appendix 2: Process Flow Diagram==
[[File:Appendix_2.png|center]]
==Appendix 3: Mass and Energy Balance Tables==
'''Table 1'''. Mass balance on 1 stage 5 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T1.png|center|500px|]]

'''Table 2'''. Mass balance on 1 stage 1 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T2.png|center|500px|]]

'''Table 3'''. Mass balance on 2 stage 5,5 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T3.png|center|600px|]]

'''Table 4'''. 1 stage 5 elements CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T4.png|center|300px|]]

'''Table 5'''. 1 stage 1 element CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T5.png|center|300px|]]

'''Table 6'''. 2 stages 5,5 elements CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T6.png|center|300px|]]

'''Table 7'''. Energy balance on 1 stage 5 elements RO system
[[File:Appendix_3_T7.png|center|800px|]]

==Appendix 4: Effluent Summary==
[[File:Appendix_4.png|center]]

==Appendix 5: Cost of Production==
[[File:Appendix_5.png|center]]

==Appendix 6: Economic Anlaysis==
===Case I ===
Water is $700; RO with 1 stage and 5 elements
[[File:Appendix_6_C1.png|center]]

===Case II===
Water is $1400; RO with 1 stage and 5 elements
[[File:Appendix_6_C2.png|center]]

===Case III: Carlsbad ===
Water is $2257; RO with 1 stage and 5 elements
[[File:Appendix_6_C3.png|center]]

===Case IV: Breakeven (NPV = 0 at 10 years)===
Water is $2104; RO with 1 stage and 5 elements
[[File:Appendix_6_C4.png|center]]

===Case V: Breakeven (NPV = 0 at 10 years)===
Water is $2112; RO with 2 stages and 5,5 elements
[[File:Appendix_6_C5.png|center]]

===Case VI: Breakeven (NPV = 0 at 10 years)===
Water is $2138 with 1 stage and 1 element
[[File:Appendix_6_C6.png|center]]

==Appendix 7: ROSA Simulation Outputs==

'''Table 1'''. System detail of 1 stage 5 elements RO system
[[File:Appendix_7_T1.png|center|500px|]]

'''Table 2'''. System detail of 1 stage 1 elements RO system
[[File:Appendix_7_T2.png|center|500px|]]

'''Table 3'''. System detail of 2 stage 5,5 elements system
[[File:Appendix_7_T3.png|center|500px|]]

'''Table 4'''. 1 stage 5 elements RO system concentrations and pH in each stream
[[File:Appendix_7_T4.png|center|500px|]]

'''Table 5'''. 1 stage 1 element RO system concentration and pH in each stream
[[File:Appendix_7_T5.png|center|500px|]]

'''Table 6'''. 2 stage 5,5 elements RO system concentrations and pH in each stream
[[File:Appendix_7_T6.png|center|700px|]]

=References=
<references/>

Desalination - Team E

2016-03-12T03:40:31Z

Bts511: /* Design Basis and Technical Approach */

Group E Corporation

Authors: Hassan Ali, Woo Soo Choe, Brett Sleyster, Jake Stolley

Instructors: Fengqi You, David Wegerer

March 11, 2016

=Executive Summary=

Outlined in this report is the project model report for the proposed desalination plant in Los Angeles, California which will provide drinking water to local citizens using reverse osmosis process. This location was chosen due to the region’s propensity for droughts in recent years. The fresh water resources in California have plummeted as water prices have risen 3-4% each year. This is only expected to continue. Other companies have worked deals with the Californian government to create desalination plants that provide a new fresh water drinking source. This suggests that there is a strong market here with potential growth.

The plant will have an output of 15 million gallons per day, which is about medium size relative to other plants. Thus the build time will be about one year. The plant will run for 30 years, with 350 days per year of use. Shutdown time will be used for cleaning and maintenance of equipment. The plant will be built on Long Beach, 500 meters from the coast nearby the Long Beach Water Reclamation Plant (LBWP). The input will withdraw 70.5 million gallons per day of seawater that will be used for reverse osmosis and dilution of brine.

Reverse osmosis was chosen as the desalination method for its low energy cost and flexibility. We designed a process that will pretreat the seawater to remove waste, colloids, bacteria, and membrane-damaging scalants. Once this seawater is treated is pumped to heavy pressures through a reverse osmosis system. After doing process optimization under various system parameter changes, we found a continuous, single stage, five element reverse osmosis process to be best. It produces a near 30% recovery. We reduce energy costs by including a pressure exchanger between our concentrate outlet and adjusted feed into the reverse osmosis system.

After performing mass and energy balances, we were able to size equipment and find our permeate composition. It turns out our process will meet L.A. standards for drinking water quality. However, post treatment will be done by the local LBWP to make the water taste better and be non-caustic to pipes. This will reduce “in-house” costs. Many environmental considerations were made, to meet our environmental standards. We meet EPA salinity requirements for our brine waste by diluting it. We also don’t use a lot of energy and thus reduce our fossil fuel consumption and CO2 generation. To be safe, we work well under max pressure limits for our equipment. We make sure we are ethical and meet all standards that our local county requires of us.

After an economic evaluation at a conservative well-water price of $700 per acre-foot, we realize that our revenue is way too low to cover our costs. Capital costs are especially large with regards to our large water tanks. We find that we reach a 10 year NPV breakpoint of $2104 per acre-foot after performing a sensitivity analysis. The Carlsbad Plant in San Diego worked out a deal for $2257 per acre-foot.

As we have met all design constraints for this project, we hope to convince policymakers to adopt our project at least $2104 per acre-foot water. If this deal is reached, we recommend that this project and design be funded. There is future work to be done to improve costs such as performing pump optimization. However, this work is minimal and won’t harm our bottom line if we reach our proposed price points.

__TOC__

=Introduction=

Due to a combination of various factors--including rising human population, climate change, increased energy consumption, and land erosion--the availability of fresh water as a resource for human consumption has been dwindling. This has had a negative impact on the world, but this impact has been especially felt in areas with natural shortages in fresh water. Desert regions, such as Saudi Arabia, have resorted to desalinating seawater at the coast and transporting up to their cities. This encompasses their entire freshwater system as very few natural water sources remain. In the United States there is a region which has shown increasing signs of following in the footsteps of Saudi Arabia and becoming fully reliant on desalination: Los Angeles, California.

=Location and Market Analysis=

In this project, Los Angeles was chosen as the location our plant, because the area has been subject to a severe drought for the last four years, as shown in Figure 1. The rapid decrease in the available water has caused permanent damages to the ground, decreased the amount of water in the reservoir from 45% to 25%, and reduced the amount of hydroelectric power produced in the state. Such deficiency in usable water has increased the water prices in Los Angeles by 6-7% annually for the last decade and is expected to rise by 3.4% annually for typical customers for the next five years. For customers who use more than 20000 gallons a year, the bill is expected to increase by 34% by 2021. This price increase is occurring because of both drought and maintenance. <ref>Pacific Institute: The California Drought. 2014.</ref>

[[File:Figure_1.jpg|right|200px|frame|Figure 1. California Drought Index.]]
As of March 2015 the Metropolitan Water District of Southern California was willing to pay as much as $700 per acre-foot of water to farmers in the Sacramento Valley, so we assume this will be the target price of the water produced by the this desalination plant.<ref>Hessel, Phil. “Dry Southern California Offers Northern Farmers Top Dollar for Water”. NBC News. March 18, 2015. </ref>

Currently, there is a plant in Carlsbad which supplies drinking water to the San Diego area. The plant can produce up to 50 million gallons per day and is the nation's largest water desalination plant. The water from this plant costs $2257 per acre-foot for the first 48000 acre-feet and about $2000 for any additional acre-foot. The county signed a contract with this plant to purchase the first 48000 acre-feet of water for the next 30 years.<ref>Carlsbad Desalination Plant. “Project Overview”. January 2016.</ref> Also, there is another desalination plant in Tampa that produces about 25 million gallons of water per day, and over 2000 other public and industrial plants with production capacity of over 300000 gallons per day. When conducting the economic analysis our plant, such numbers presented by other plants must be considered, so our project can yield positive profit.<ref>Leven, Rachel. “U.S. Desalination Industry Grows Since 2000; See as Essential to Meeting Supply Needs”. Bloomberg BNA. August 21, 2013.</ref>

The profitability of the project is naturally heavily dependent on the price of water in California. If the drought ends, water supply will increase and the water price will drop significantly, but if the drought continues, the price of water in California could rise even more rapidly and allow desalination plants to be more profitable.<ref>Southern California Public Radio. “Where is California water use decreasing?”. December 2015.</ref> It may be difficult to make this plant profitable without a guaranteed purchase or other subsidy from the local governments because the presence of pre-existing plants maintains the water price relatively low. However, external sources of profit such as guaranteed purchase or subsidy are expected because the ramifications of not having desalination plant or other sources of water may be significantly more expensive than the subsidizing a desalination plant assuming the drought continues.<ref>Weiser, Matt. “Could desalination solve California’s water problem?”. The Sacramento Bee. October 18, 2014.</ref>

=Design Basis and Technical Approach=

[[File:Figure_2.jpg|left|200px|frame|Figure 2. Vertical Well Feed]]
Considering the number of pre-existing desalination plant and the potential risk of the drought ending, our plant will be tentatively designed to produce 15 million gallons of drinkable water per day to homes and businesses in the Los Angeles county area. With a recovery rate of near 30%, and a requirement for diluting any brine, the input feed flowrate will be 70.5 million gallons per day. Following the similarly built Carlsbad plant, we expect the plant to last 30 years. The project building timeline is one year based on our relatively small plant size. We will run for 350 days a year while shutting down for cleaning and maintenance once every 3 months (based on DOW standards).<ref name = "Bates">Bates, Wayne. “Cleaning your RO”. Hydranautics. </ref> The plant will be built 500 meters from the coastline on Long Beach. A borehole will be drilled 10 meters into the ground. At this depth there should be seawater aquifers below the coast that regularly recharges itself with seawater from the ocean. A vertical well containing intake holes will be put into the borehole. The intake holes will be equipped with large size filters that can easily be replaced. At the surface, the well will be connected to a large-scale pump in a well pump house. From there, the seawater will be horizontally pumped into a seawater tank in the desalination plant. This will be considered the start of our process. The process ends after pretreatment, desalination, post treatment, and brine treatment are completed. Appendix 1 shows more details on the Design Basis.

The feed composition is assumed to be the standard seawater “average” based on the DOW Chemical ROSA 9.0 Technical Manual in Table 1. However, in reality the seawater composition at the plant site will have to be chemically tested. The final product will be based on L.A. standards for drinking water quality (Table 1). The specifications in the table are upper limits that we cannot exceed. We will aim to reach the salt level standards but any remineralization and pH treatment will be exported and paid for due to the added complexity and cost in making such a system in-house.

Due to the acidity of the desalinated water coming out of the reverse osmosis (RO) system (pH ~ 6), the pumps and pipes carrying the desalinated water for post treatment will likely corrode over time. We have decided to place the desalination plant by the Long Beach Water Reclamation Plant (about 50 meters away) to prevent pipe corrosion.

[[File:Table_1.jpg|center|800px|frame|Table 1. Concentration of the standard seawater, the output from the RO system, and LA standard.]]

=Desalination Overview and Alternatives=

We first began our research into the current landscape of the desalination industry, particularly of plants that were built in the United States. We discovered there were many processes but the main ones included: 
* Membrane Separation (Forward Osmosis and Reverse Osmosis)
* Thermal Separation (Multi Stage Flash Distillation)
* Phase Change Separation (Freeze Separation)
Each process was analyzed for their pros and cons. For our project design, we decided to go with reverse osmosis as our desalination process. This is because it is the most commonly used process in the industry, and for good reason. Unlike thermal separation, which requires huge amounts of energy to supply heat, reverse osmosis is very energy efficient. This reduces energy costs that would be spread across the plant’s lifetime.<ref name = "Republic">Republic of Mauritius NY. "Technical Aspects of Desalination Plant". Sustainable Sanitation and Water Management. </ref>

Forward osmosis is a newer technology being developed that also uses membranes to separate water and salt and may be more energy efficient overall than reverse osmosis. This is because forward osmosis doesn’t require a pump to pressurize seawater through the membranes, but rather leaves a higher salt concentrated solution (AKA draw solution) on the other side of the membrane for the seawater to be drawn into. Then a second step is required to regenerate the concentrated draw solution and produce purified water.<ref>Mallinson, Alissa. "Study Shows Forward Osmosis Desalination Not Energy Efficient." MIT News [Boston]. July 23, 2014.</ref> We decided to move away from forward osmosis because it is a process that is still being refined and developed for desalination and lacks the data, programs, and technologies that reverse osmosis users benefit from. Additionally, a recent study from MIT has claimed that due to the inherent nature of forward osmosis requiring that seawater be drawn to a higher concentrated salt solution before the water and salt is separated, it cannot possibly be more energy efficient than reverse osmosis.<ref>O'Hern, Sean C., Karnik, Rohit. Nanofiltration across Defect-Sealed Nanoporous Monolayer Graphene. American Chemical Society 15.5 (2014); 3254-260. </ref>

Other processes such as freeze separation were too restrictive in their feed and process demands and required a lot of capital. Ultimately, reverse osmosis is simply the best process that exists for desalination and has been ranked the best when performance is evaluated in terms of technical, environmental, and economic aspects. It is a process that is flexible in water quantity, operation start-up and shut-off, and site location.

Nevertheless there are disadvantages to reverse osmosis that must be considered. The feed seawater must not be too salty and too low quality. The pressure to pump the seawater is enormous. The construction time is long for large plants. The capital costs are also high when compared with other processes such as thermal separation. But all of these disadvantages are ultimately outweighed by the advantages reverse osmosis provides.<ref name = "Republic"/>

=Process Overview and Alternatives=

See Appendix 2 (Figure 1) for the Process Flow Diagram (PFD). The process itself is detailed below. The stream flow rates are in Appendix 2 (Table 1). We assumed 10 meters of pipe for every stream except the inlet and outlet streams of the entire PFD.

==Pretreatment==

After the 10 meter vertical well pumps seawater from the Long Beach coast, the feed is pumped 500 meters to a seawater tank that makes it easy to shut-down the system and reduce the large volume of flow. The level control keeps the seawater from overfilling the tank. From there, part of the seawater goes to brine treatment while the rest go to pretreatment. The purpose of pretreatment in the desalination process is to prevent equipment damage by removing particles and microorganisms from the water. A variety of pollutants -- including calcium carbonate, sulfates, silica, clay, bacteria, and other colloids and microorganisms -- can disrupt the reverse osmosis process. Media filtration uses fine grains such as sand or anthracite to filter the particles. We used a large filter to catch larger particle material (e.g. sand). However, simple filtration is not enough to prevent colloid fouling. An ultrafilter (i.e. microfiltration) is used to remove virtually all suspended matter, effectively preventing colloid fouling. The costs associated with ultrafiltration are not prohibitive in terms of the project’s budget.<ref name = "Desalination">Desalination Pre-treatment. Lenntech.</ref>

Reverse osmosis systems can also be disrupted by the growth of microbiological films of bacteria on the film and pipe walls. They are typically treated as colloids during pretreatment, and most organisms are removed from the water by ultrafiltration, but this does not remove the smallest of organic particles. DBNPA is the most effective biocide, but these aren’t often used and have to be removed from the water. Chlorination is an effective way to remove bacteria, but it can cause degradation of the membrane, and must be removed before the reverse osmosis occurs. In our process we will use UV irradiation to remove any organic material that passes through the microfilters.<ref name = "Desalination"/>

Scaling can reduce the flux across the membrane and can build up in pipe corners, causing backup. It is caused by inorganic salts -- specifically sulfates and calcium and magnesium carbonate -- present in the water. Scaling can be prevented and/or controlled with acids, antiscalants, or ion exchange softening. Antiscalants are good for removing sulfates, but not much else. Acidification will only be effective at removing CaCO3 from water but, the decreased pH may deteriorate the membrane and pipes, so it may require future research.<ref name = "Desalination"/> Ion exchange is very expensive and requires more capital to undergo, but is overall the more effective option.

The ion exchanger is a pretreatment system which acts as a method to fight membrane scaling in reverse osmosis. The ion exchanger works by using sodium carboxylate as a resin which will react with Chlorine and Magnesium ions, removing them from the feed. The sodium is substituted out and enters the feed stream, increasing the overall sodium level and TDS. The lack of magnesium and chlorine ions means that there is nothing to react with the carbonates to form scaling salts on the membranes.

==Desalination (Reverse Osmosis)==

Two pumps (P-105 & P-106) take the adjusted feed seawater up to 58 bar. However, before this pumping occurs, about 70% of the original stream gets separated and heads into a PX Pressure Exchanger. This separated stream later joins together with the pumped stream to enter the reverse osmosis system at the required pressure of 58 bar. The streams can combine together as they have equal pressure. It is the concentrate stream from the output of the reverse osmosis system that transfers its energy to the adjusted feed in the PX Pressure Exchanger.

[[File:Figure_3.jpg|left|300px|frame|Figure 3. Inside look of pressure exchanger]]
The pressure of the concentrate stream is about 57 bar and it is fed into an exchanger which works by using vessel that resembles a small pipe within a rotating wheel. The low pressure (LP) feed enters the vessel at 1 bar (see Figure 3). The feed turns and is pushed out into an outlet pipe by the high pressure (HP) concentrate stream at ~57 bar. The feed is now a HP stream via this exchanged energy. The now LP concentrate stream turns and is pushed out into an outlet pipe by the LP feed entering the vessel.

98% of the energy within the concentrate stream can be recovered via the PX Pressure Exchanger.<ref>Desalination Product and Applications Catalog 2013. “Advance Technologies for Seawater Reverse Osmosis and Brackish Water Applications”. Energy Recovery, Inc. 2012-2014.</ref >Thus, the now HP feed stream is not quite 58 bar and will need to be pumped by a small booster pump (P-107).

When designing a reverse osmosis system, we need to make decisions on what specific kind of reverse osmosis system will be implemented. When making the decision, we investigated the advantages of different types of reverse osmosis among the following options: batch vs. continuous process, number of stages, and number of elements.<ref>FILMTEC Reverse Osmosis Membranes. “Water & Process Solutions Technical Manual.” Dow Chemical Company.</ref> Our decision was made based on process optimization which is described later in this report. The end result of that optimization was to go with a single stage, continuous process with five elements in series.

=Post Treatment=

The permeate is sent to post treatment after the salt content is analyzed and controlled if too high or too low due to any unexpected errors. As stated above, the post treatment process will be done 50 meters from our plant by the Long Beach Water Reclamation Plant. The reason we aren’t performing post treatment “in-house” is due to the increased cost it would take for us to build and run the equipment for post treatment. We have the option to use the government’s water treatment facilities that already exist and our 15 million gallons per day output is a relatively small addition to the large Long Beach Water Reclamation Plant. We will assume any costs to post treatment will be pushed to the consumer in the price of purchasing the drinking water from the Long Beach Water Reclamation Plant.

Post treatment at any facility works to bring water to “taste standards”. The water that is the permeate of our process is already meeting L.A. drinking standards as shown above in Table 1 after our mass balance calculations around the reverse osmosis system (Appendix 3, Table 1). The post treatment increases the pH from 6.3 to 7.3 with the addition of base and then adds minerals like fluoride, calcium, and magnesium back in to improve taste.

==Brine Treatment and Byproduct Waste==

A natural consequence of a reverse osmosis procedure is that a lot of salt will be leftover from the membrane separation in a very concentrated form. This is called “brine” and will have to be taken care of. According to EPA requirements, the salt concentration of water entering the environment (via land or sea) cannot be higher than 40,000 ppm.8 The salt concentration of our brine is much higher (around 70,000 ppm). Thus, the brine must be removed or treated. We decided to treat the brine by diluting it as removing brine can be difficult and may require lots of land (e.g. evaporation ponds).

After the concentrate exchanges energy in the energy exchanger, it becomes a low pressure stream. It is pumped into a dilution tank and mixes with 18.2 million gallons per day of seawater from the inlet to become diluted brine which is pumped 600 meters out to sea.

The waste byproducts are the solid buildup on the filters and membranes. The filters and membranes also have a set lifetime and thus will also need to be removed. The waste is all non-hazardous and thus can be put in the municipal dump.

=Process Details and Equipment Sizing=

==Filters and UV Treatment [Pretreatment]==
The cost of our ultrafiltration system is estimated using data from the University of Tennessee at Knoxville.<ref>Qiang He and R. Bruce Robinson Membrane Filtration Systems. “Membrane Module Costs.” University of Tennessee at Knoxville. December 2014.</ref> It estimated the cost of the cost of the ultrafiltration $0.057 per 1000 gallons and the construction and capital costs are $0.39 per gallon per day. The cost of UV treatment is based off of units available from buyultraviolet.com for units that can treat about 450 gallons per minute.<ref>Atlantic Ultraviolet Corporation. “Megatron Water Disinfection System.”</ref> Due to the water having to be in relatively small pipes for treatment we do not expect the costs to scale very much, so we believe these units to be an accurate cost estimate.

==Ion Exchanger [Pretreatment] ==

The ion exchanger was sized using data from DOW. Our process will use 60 vessels with ion exchange beds of volume 5.73 m3. With a flow rate of 30 bed volumes/hour (10,312.2 m3/hr) the system of beds has a capacity of 125% of the feed. This allows for periodic regeneration of the resin, a weak acid cation resin (DOW MAC-3 Resin). This weak cation resin will remove cationic calcium and magnesium from the water to prevent scaling. The mass flow was calculated via a mass balance on the ion exchanger (Appendix 3).

==Reverse Osmosis System [Desalination]==

Each of the 2607 Reverse Osmosis SW30HRLE-440i [Seawater 30 High Rejection Low Energy 440 ft2 interlocking cap membrane] units have the same dimensions. A, the net length, is 1016 mm. B, the total length including the endcap, is 1029 mm. C, the diameter of the filter element, is 201mm. D, the diameter of the permeate pathway, is 29 mm. For a single stage, we have five of the elements in series so the overall length of a single stage is expected to be roughly 5145 mm. The membrane material is polyamide thin film composite. It will last three years.
Operating Conditions [ROSA 9.0]: 
* Temperature: 30oC 
* Pressure: 58 bar 
* Flow Rate: 52.3 Million gal per day 
* Recovery: 28.7%

[[File:Figure_4.png|center|400px|frame|Figure 4. Diagram of Dow FILMTEC SW30HRLE-440i]]
Cleaning Schedule: 
* Conservative Estimate: Once every three months (three days available for cleaning)<ref name = "Bates"/> 
* DOW FILMTEC Cleaning Conditions 
** The normalized permeate flow drops 10%
* The normalized salt passage increases 5 - 10%

==Seawater Holding Tank and Dilution Tank [Brine Treatment]==

The initial seawater holding tank works as a buffer to prevent surges and help perform shutdowns quickly and easily. We expect to run continuously for 350 days a year with the rest of the available time spread out and used for shutting down the plant for cleaning membranes, filters, and the fixing of unexpected errors. There is a big tradeoff between having a holding tank with high retention time and having a holding tank with low cost. With that in mind, we erred on the side of lowering capital cost and went with a retention time of one hour. The volume could be found by multiplying the retention time with the feed flow rate (Stream 1). We assumed that a good height to width ratio for a large water tank would be 4:1.

A mass balance was done around the dilution tank to find what amount of seawater would be necessary to dilute the brine to 45,000 ppm salinity (Appendix 3; Table 3). With this the feed flow rate into the tank could be found. It was found that it takes about 20 minutes of retention time for good dilution of the brine in seawater.8 We again assumed that a good height to width ratio for a large water tank would be 4:1.

We assumed a corrosion allowance of 4 mm wall thickness. We calculated wall thickness using vessel pressure, hoop stress, longitudinal stress, and weld factor. Each tank is made with SS316 to protect against corrosion.<ref>Towler G, Sinnott R. Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013.</ref>

<center>
'''Table 2.''' Tank sizes.
[[File:Table_2.jpg|center|600px]]
</center>

=Process Optimization=
<center> '''Table 3.'''Optimization Criteria
[[File:Table_3.jpg|center|400px]]
</center>
In order to optimize the reverse osmosis system, numerous simulations were run. Three representative cases are reported as shown above (Table 3). The objective of the optimization was to minimize the required water price to break even in 10 years. In order to achieve this goal, the balance between the capital cost from the number of the filter elements and utility cost from energy consumption was weighed. Before the energy recovery system was implemented, the general trend was that if we have more filter elements in series, overall recovery increases but recovery rate per filter decreases, and we need more filter elements. While increasing the elements in series increase the required number of filters, and thereby increases the capital cost, we require less energy to produce the same amount of water. After the energy recovery system is implemented, the correlation between the number of elements in series and the required energy becomes more complicated because high pressure concentrate is used to pressurize the feed. However, the trade-off between the utility cost and capital cost still exists. Of the three cases shown above, single stage, five elements RO system turned out to be the most optimal solution because single stage, single element system required higher utility cost and two stage system slightly did not meet the LA standard. Detailed information which was used for the optimization is available in Appendix 3 and in Appendix 6.

=Ethical, Environmental, and Societal Considerations=

There are a few key considerations we have followed to meet ethical, environmental, and societal guidelines and requirements: 
* Meet L.A. Drinking Water Standards
** Correct ion concentration
* Moved near water treatment plant
** Pipes will not corrode from low pH
* Working below pressure limits
* Water will be locally served and meet taste standards
** Via Post Treatment mineralization
* Follows EPA Environmental Salinity Requirements
* Reduced Energy Requirements
** PX Pressure Exchanger
** Optimized RO System
** Reverse Osmosis is itself Energy Saving

For further environmental considerations, see effluent summary in Appendix 4.

=Economic Analysis=

The pricing of the majority of the process was completed using the Aspen Economic Analyzer, however we had determine costs for the ultrafiltration, reverse osmosis membranes, and the ultraviolet treatment was completed separately. This led to a total ISBL of $56 MM, OSBL of $22.4 MM, Engineering Costs of $22.4 MM and a Contingency of $7.8 MM. We also chose a working Capital of $10 MM. This working capital is sufficient because we have no raw materials costs. We also had FCOP of $5.2 MM and VCOP of $3.5 MM.

We originally were planning on pricing our water at $700 per acre-foot, but after doing economic analysis we had a 10 year NPV of $-76.2 MM. After determining did a sensitivity analysis and determined that if our water was priced at $2104 per acre-foot we could have an NPV of $0 after 10 years and if we were able to get the same price of $2257 per acre-foot as the plant in Carlsbad we would have a 10 year NPV of $8.3 MM which increases to $41.1 MM at 20 years. For more detailed data about the economic analysis see Appendix 6.

=Conclusion and Recommendations=

Our mission was to create an economically viable process that will desalinate water for human consumption in the L.A. County. We have found after research and optimization that the best process to undergo desalination is a continuous single stage, five elements reverse osmosis system. Our process can meet the required L.A. drinking water standard and we have set up plans for post treatment at a local water reclamation site. While meeting the drinking water standards, we also met our ethical, environmental, and social standards.

Nevertheless, at a sale price of $700 per acre-foot, our proposed desalination plant in Los Angeles, California looks to be a very expensive project to undergo. Without any sort of government assistance to carry the heavy capital costs of such a project, we would suggest the plant not be built. However, after performing a sensitivity analysis at higher (and less conservative) price points we reached a breakeven point in 10 years with a sale price of $2104 per acre-foot of drinking water. Research has shown that the droughts striking California look only to get worse in the near future. With limited freshwater resources, there is no option other than to consider desalination. Therefore it is reasonable to expect we could get near the same price deal as local desalination plants like the Carlsbad Plant ($2257 per acre-foot) As we have met all design constraints for this project, we hope to convince policymakers to adopt our project at least $2104 per acre-foot water. If this deal is reached, we recommend that this project and design be funded.

However, there are ways to improve our design with future work. Right now all the pipes diameters are based on having a flowrate of 1.5 m/s. But in reality, we’d need to use standard piping sizes to get more accurate pricing and energy costs. Also, in reality we would be purchasing some of these vessels in bulk quantities (i.e. bulk filters) so the price could be lowered with negotiation and economies of scale. We also have pumps that are providing very little pressure for their high cost. We can remove these pumps and simply run at higher pressures earlier. We can potentially reduce early energy costs by building a “floating” desalination plant in the sea to reduce pumping distance from the well.

=Appendices=

==Appendix 1: Design Basis==
[[File:Appendix_1.png|center|Appendix_1.png]]

==Appendix 2: Process Flow Diagram==
[[File:Appendix_2.png|center]]
==Appendix 3: Mass and Energy Balance Tables==
'''Table 1'''. Mass balance on 1 stage 5 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T1.png|center|500px|]]

'''Table 2'''. Mass balance on 1 stage 1 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T2.png|center|500px|]]

'''Table 3'''. Mass balance on 2 stage 5,5 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T3.png|center|600px|]]

'''Table 4'''. 1 stage 5 elements CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T4.png|center|300px|]]

'''Table 5'''. 1 stage 1 element CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T5.png|center|300px|]]

'''Table 6'''. 2 stages 5,5 elements CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T6.png|center|300px|]]

'''Table 7'''. Energy balance on 1 stage 5 elements RO system
[[File:Appendix_3_T7.png|center|800px|]]

==Appendix 4: Effluent Summary==
[[File:Appendix_4.png|center]]

==Appendix 5: Cost of Production==
[[File:Appendix_5.png|center]]

==Appendix 6: Economic Anlaysis==
===Case I ===
Water is $700; RO with 1 stage and 5 elements
[[File:Appendix_6_C1.png|center]]

===Case II===
Water is $1400; RO with 1 stage and 5 elements
[[File:Appendix_6_C2.png|center]]

===Case III: Carlsbad ===
Water is $2257; RO with 1 stage and 5 elements
[[File:Appendix_6_C3.png|center]]

===Case IV: Breakeven (NPV = 0 at 10 years)===
Water is $2104; RO with 1 stage and 5 elements
[[File:Appendix_6_C4.png|center]]

===Case V: Breakeven (NPV = 0 at 10 years)===
Water is $2112; RO with 2 stages and 5,5 elements
[[File:Appendix_6_C5.png|center]]

===Case VI: Breakeven (NPV = 0 at 10 years)===
Water is $2138 with 1 stage and 1 element
[[File:Appendix_6_C6.png|center]]

==Appendix 7: ROSA Simulation Outputs==

'''Table 1'''. System detail of 1 stage 5 elements RO system
[[File:Appendix_7_T1.png|center|500px|]]

'''Table 2'''. System detail of 1 stage 1 elements RO system
[[File:Appendix_7_T2.png|center|500px|]]

'''Table 3'''. System detail of 2 stage 5,5 elements system
[[File:Appendix_7_T3.png|center|500px|]]

'''Table 4'''. 1 stage 5 elements RO system concentrations and pH in each stream
[[File:Appendix_7_T4.png|center|500px|]]

'''Table 5'''. 1 stage 1 element RO system concentration and pH in each stream
[[File:Appendix_7_T5.png|center|500px|]]

'''Table 6'''. 2 stage 5,5 elements RO system concentrations and pH in each stream
[[File:Appendix_7_T6.png|center|700px|]]

=References=
<references/>

Desalination - Team E

2016-03-12T03:39:53Z

Bts511: /* Design Basis and Technical Approach */

Group E Corporation

Authors: Hassan Ali, Woo Soo Choe, Brett Sleyster, Jake Stolley

Instructors: Fengqi You, David Wegerer

March 11, 2016

=Executive Summary=

Outlined in this report is the project model report for the proposed desalination plant in Los Angeles, California which will provide drinking water to local citizens using reverse osmosis process. This location was chosen due to the region’s propensity for droughts in recent years. The fresh water resources in California have plummeted as water prices have risen 3-4% each year. This is only expected to continue. Other companies have worked deals with the Californian government to create desalination plants that provide a new fresh water drinking source. This suggests that there is a strong market here with potential growth.

The plant will have an output of 15 million gallons per day, which is about medium size relative to other plants. Thus the build time will be about one year. The plant will run for 30 years, with 350 days per year of use. Shutdown time will be used for cleaning and maintenance of equipment. The plant will be built on Long Beach, 500 meters from the coast nearby the Long Beach Water Reclamation Plant (LBWP). The input will withdraw 70.5 million gallons per day of seawater that will be used for reverse osmosis and dilution of brine.

Reverse osmosis was chosen as the desalination method for its low energy cost and flexibility. We designed a process that will pretreat the seawater to remove waste, colloids, bacteria, and membrane-damaging scalants. Once this seawater is treated is pumped to heavy pressures through a reverse osmosis system. After doing process optimization under various system parameter changes, we found a continuous, single stage, five element reverse osmosis process to be best. It produces a near 30% recovery. We reduce energy costs by including a pressure exchanger between our concentrate outlet and adjusted feed into the reverse osmosis system.

After performing mass and energy balances, we were able to size equipment and find our permeate composition. It turns out our process will meet L.A. standards for drinking water quality. However, post treatment will be done by the local LBWP to make the water taste better and be non-caustic to pipes. This will reduce “in-house” costs. Many environmental considerations were made, to meet our environmental standards. We meet EPA salinity requirements for our brine waste by diluting it. We also don’t use a lot of energy and thus reduce our fossil fuel consumption and CO2 generation. To be safe, we work well under max pressure limits for our equipment. We make sure we are ethical and meet all standards that our local county requires of us.

After an economic evaluation at a conservative well-water price of $700 per acre-foot, we realize that our revenue is way too low to cover our costs. Capital costs are especially large with regards to our large water tanks. We find that we reach a 10 year NPV breakpoint of $2104 per acre-foot after performing a sensitivity analysis. The Carlsbad Plant in San Diego worked out a deal for $2257 per acre-foot.

As we have met all design constraints for this project, we hope to convince policymakers to adopt our project at least $2104 per acre-foot water. If this deal is reached, we recommend that this project and design be funded. There is future work to be done to improve costs such as performing pump optimization. However, this work is minimal and won’t harm our bottom line if we reach our proposed price points.

__TOC__

=Introduction=

Due to a combination of various factors--including rising human population, climate change, increased energy consumption, and land erosion--the availability of fresh water as a resource for human consumption has been dwindling. This has had a negative impact on the world, but this impact has been especially felt in areas with natural shortages in fresh water. Desert regions, such as Saudi Arabia, have resorted to desalinating seawater at the coast and transporting up to their cities. This encompasses their entire freshwater system as very few natural water sources remain. In the United States there is a region which has shown increasing signs of following in the footsteps of Saudi Arabia and becoming fully reliant on desalination: Los Angeles, California.

=Location and Market Analysis=

In this project, Los Angeles was chosen as the location our plant, because the area has been subject to a severe drought for the last four years, as shown in Figure 1. The rapid decrease in the available water has caused permanent damages to the ground, decreased the amount of water in the reservoir from 45% to 25%, and reduced the amount of hydroelectric power produced in the state. Such deficiency in usable water has increased the water prices in Los Angeles by 6-7% annually for the last decade and is expected to rise by 3.4% annually for typical customers for the next five years. For customers who use more than 20000 gallons a year, the bill is expected to increase by 34% by 2021. This price increase is occurring because of both drought and maintenance. <ref>Pacific Institute: The California Drought. 2014.</ref>

[[File:Figure_1.jpg|right|200px|frame|Figure 1. California Drought Index.]]
As of March 2015 the Metropolitan Water District of Southern California was willing to pay as much as $700 per acre-foot of water to farmers in the Sacramento Valley, so we assume this will be the target price of the water produced by the this desalination plant.<ref>Hessel, Phil. “Dry Southern California Offers Northern Farmers Top Dollar for Water”. NBC News. March 18, 2015. </ref>

Currently, there is a plant in Carlsbad which supplies drinking water to the San Diego area. The plant can produce up to 50 million gallons per day and is the nation's largest water desalination plant. The water from this plant costs $2257 per acre-foot for the first 48000 acre-feet and about $2000 for any additional acre-foot. The county signed a contract with this plant to purchase the first 48000 acre-feet of water for the next 30 years.<ref>Carlsbad Desalination Plant. “Project Overview”. January 2016.</ref> Also, there is another desalination plant in Tampa that produces about 25 million gallons of water per day, and over 2000 other public and industrial plants with production capacity of over 300000 gallons per day. When conducting the economic analysis our plant, such numbers presented by other plants must be considered, so our project can yield positive profit.<ref>Leven, Rachel. “U.S. Desalination Industry Grows Since 2000; See as Essential to Meeting Supply Needs”. Bloomberg BNA. August 21, 2013.</ref>

The profitability of the project is naturally heavily dependent on the price of water in California. If the drought ends, water supply will increase and the water price will drop significantly, but if the drought continues, the price of water in California could rise even more rapidly and allow desalination plants to be more profitable.<ref>Southern California Public Radio. “Where is California water use decreasing?”. December 2015.</ref> It may be difficult to make this plant profitable without a guaranteed purchase or other subsidy from the local governments because the presence of pre-existing plants maintains the water price relatively low. However, external sources of profit such as guaranteed purchase or subsidy are expected because the ramifications of not having desalination plant or other sources of water may be significantly more expensive than the subsidizing a desalination plant assuming the drought continues.<ref>Weiser, Matt. “Could desalination solve California’s water problem?”. The Sacramento Bee. October 18, 2014.</ref>

=Design Basis and Technical Approach=

[[File:Figure_2.jpg|left|200px|frame|Figure 2. Vertical Well Feed]]
Considering the number of pre-existing desalination plant and the potential risk of the drought ending, our plant will be tentatively designed to produce 15 million gallons of drinkable water per day to homes and businesses in the Los Angeles county area. With a recovery rate of near 30%, and a requirement for diluting any brine, the input feed flowrate will be 70.5 million gallons per day. Following the similarly built Carlsbad plant, we expect the plant to last 30 years. The project building timeline is one year based on our relatively small plant size. We will run for 350 days a year while shutting down for cleaning and maintenance once every 3 months (based on DOW standards).<ref name = "Bates">Bates, Wayne. “Cleaning your RO”. Hydranautics. </ref> The plant will be built 500 meters from the coastline on Long Beach. A borehole will be drilled 10 meters into the ground. At this depth there should be seawater aquifers below the coast that regularly recharges itself with seawater from the ocean. A vertical well containing intake holes will be put into the borehole. The intake holes will be equipped with large size filters that can easily be replaced. At the surface, the well will be connected to a large-scale pump in a well pump house. From there, the seawater will be horizontally pumped into a seawater tank in the desalination plant. This will be considered the start of our process. The process ends after pretreatment, desalination, post treatment, and brine treatment are completed. Appendix 1 shows more details on the Design Basis.

The feed composition is assumed to be the standard seawater “average” based on the DOW Chemical ROSA 9.0 Technical Manual in Table 1. However, in reality the seawater composition at the plant site will have to be chemically tested. The final product will be based on L.A. standards for drinking water quality (Table 1). The specifications in the table are upper limits that we cannot exceed. We will aim to reach the salt level standards but any remineralization and pH treatment will be exported and paid for due to the added complexity and cost in making such a system in-house.

Due to the acidity of the desalinated water coming out of the reverse osmosis (RO) system (pH ~ 6), the pumps and pipes carrying the desalinated water for post treatment will likely corrode over time. We have decided to place the desalination plant by the Long Beach Water Reclamation Plant (about 50 meters away) to prevent pipe corrosion.

[[File:Table_1.jpg|center|400px|frame|Table 1. Concentration of the standard seawater, the output from the RO system, and LA standard.]]

=Desalination Overview and Alternatives=

We first began our research into the current landscape of the desalination industry, particularly of plants that were built in the United States. We discovered there were many processes but the main ones included: 
* Membrane Separation (Forward Osmosis and Reverse Osmosis)
* Thermal Separation (Multi Stage Flash Distillation)
* Phase Change Separation (Freeze Separation)
Each process was analyzed for their pros and cons. For our project design, we decided to go with reverse osmosis as our desalination process. This is because it is the most commonly used process in the industry, and for good reason. Unlike thermal separation, which requires huge amounts of energy to supply heat, reverse osmosis is very energy efficient. This reduces energy costs that would be spread across the plant’s lifetime.<ref name = "Republic">Republic of Mauritius NY. "Technical Aspects of Desalination Plant". Sustainable Sanitation and Water Management. </ref>

Forward osmosis is a newer technology being developed that also uses membranes to separate water and salt and may be more energy efficient overall than reverse osmosis. This is because forward osmosis doesn’t require a pump to pressurize seawater through the membranes, but rather leaves a higher salt concentrated solution (AKA draw solution) on the other side of the membrane for the seawater to be drawn into. Then a second step is required to regenerate the concentrated draw solution and produce purified water.<ref>Mallinson, Alissa. "Study Shows Forward Osmosis Desalination Not Energy Efficient." MIT News [Boston]. July 23, 2014.</ref> We decided to move away from forward osmosis because it is a process that is still being refined and developed for desalination and lacks the data, programs, and technologies that reverse osmosis users benefit from. Additionally, a recent study from MIT has claimed that due to the inherent nature of forward osmosis requiring that seawater be drawn to a higher concentrated salt solution before the water and salt is separated, it cannot possibly be more energy efficient than reverse osmosis.<ref>O'Hern, Sean C., Karnik, Rohit. Nanofiltration across Defect-Sealed Nanoporous Monolayer Graphene. American Chemical Society 15.5 (2014); 3254-260. </ref>

Other processes such as freeze separation were too restrictive in their feed and process demands and required a lot of capital. Ultimately, reverse osmosis is simply the best process that exists for desalination and has been ranked the best when performance is evaluated in terms of technical, environmental, and economic aspects. It is a process that is flexible in water quantity, operation start-up and shut-off, and site location.

Nevertheless there are disadvantages to reverse osmosis that must be considered. The feed seawater must not be too salty and too low quality. The pressure to pump the seawater is enormous. The construction time is long for large plants. The capital costs are also high when compared with other processes such as thermal separation. But all of these disadvantages are ultimately outweighed by the advantages reverse osmosis provides.<ref name = "Republic"/>

=Process Overview and Alternatives=

See Appendix 2 (Figure 1) for the Process Flow Diagram (PFD). The process itself is detailed below. The stream flow rates are in Appendix 2 (Table 1). We assumed 10 meters of pipe for every stream except the inlet and outlet streams of the entire PFD.

==Pretreatment==

After the 10 meter vertical well pumps seawater from the Long Beach coast, the feed is pumped 500 meters to a seawater tank that makes it easy to shut-down the system and reduce the large volume of flow. The level control keeps the seawater from overfilling the tank. From there, part of the seawater goes to brine treatment while the rest go to pretreatment. The purpose of pretreatment in the desalination process is to prevent equipment damage by removing particles and microorganisms from the water. A variety of pollutants -- including calcium carbonate, sulfates, silica, clay, bacteria, and other colloids and microorganisms -- can disrupt the reverse osmosis process. Media filtration uses fine grains such as sand or anthracite to filter the particles. We used a large filter to catch larger particle material (e.g. sand). However, simple filtration is not enough to prevent colloid fouling. An ultrafilter (i.e. microfiltration) is used to remove virtually all suspended matter, effectively preventing colloid fouling. The costs associated with ultrafiltration are not prohibitive in terms of the project’s budget.<ref name = "Desalination">Desalination Pre-treatment. Lenntech.</ref>

Reverse osmosis systems can also be disrupted by the growth of microbiological films of bacteria on the film and pipe walls. They are typically treated as colloids during pretreatment, and most organisms are removed from the water by ultrafiltration, but this does not remove the smallest of organic particles. DBNPA is the most effective biocide, but these aren’t often used and have to be removed from the water. Chlorination is an effective way to remove bacteria, but it can cause degradation of the membrane, and must be removed before the reverse osmosis occurs. In our process we will use UV irradiation to remove any organic material that passes through the microfilters.<ref name = "Desalination"/>

Scaling can reduce the flux across the membrane and can build up in pipe corners, causing backup. It is caused by inorganic salts -- specifically sulfates and calcium and magnesium carbonate -- present in the water. Scaling can be prevented and/or controlled with acids, antiscalants, or ion exchange softening. Antiscalants are good for removing sulfates, but not much else. Acidification will only be effective at removing CaCO3 from water but, the decreased pH may deteriorate the membrane and pipes, so it may require future research.<ref name = "Desalination"/> Ion exchange is very expensive and requires more capital to undergo, but is overall the more effective option.

The ion exchanger is a pretreatment system which acts as a method to fight membrane scaling in reverse osmosis. The ion exchanger works by using sodium carboxylate as a resin which will react with Chlorine and Magnesium ions, removing them from the feed. The sodium is substituted out and enters the feed stream, increasing the overall sodium level and TDS. The lack of magnesium and chlorine ions means that there is nothing to react with the carbonates to form scaling salts on the membranes.

==Desalination (Reverse Osmosis)==

Two pumps (P-105 & P-106) take the adjusted feed seawater up to 58 bar. However, before this pumping occurs, about 70% of the original stream gets separated and heads into a PX Pressure Exchanger. This separated stream later joins together with the pumped stream to enter the reverse osmosis system at the required pressure of 58 bar. The streams can combine together as they have equal pressure. It is the concentrate stream from the output of the reverse osmosis system that transfers its energy to the adjusted feed in the PX Pressure Exchanger.

[[File:Figure_3.jpg|left|300px|frame|Figure 3. Inside look of pressure exchanger]]
The pressure of the concentrate stream is about 57 bar and it is fed into an exchanger which works by using vessel that resembles a small pipe within a rotating wheel. The low pressure (LP) feed enters the vessel at 1 bar (see Figure 3). The feed turns and is pushed out into an outlet pipe by the high pressure (HP) concentrate stream at ~57 bar. The feed is now a HP stream via this exchanged energy. The now LP concentrate stream turns and is pushed out into an outlet pipe by the LP feed entering the vessel.

98% of the energy within the concentrate stream can be recovered via the PX Pressure Exchanger.<ref>Desalination Product and Applications Catalog 2013. “Advance Technologies for Seawater Reverse Osmosis and Brackish Water Applications”. Energy Recovery, Inc. 2012-2014.</ref >Thus, the now HP feed stream is not quite 58 bar and will need to be pumped by a small booster pump (P-107).

When designing a reverse osmosis system, we need to make decisions on what specific kind of reverse osmosis system will be implemented. When making the decision, we investigated the advantages of different types of reverse osmosis among the following options: batch vs. continuous process, number of stages, and number of elements.<ref>FILMTEC Reverse Osmosis Membranes. “Water & Process Solutions Technical Manual.” Dow Chemical Company.</ref> Our decision was made based on process optimization which is described later in this report. The end result of that optimization was to go with a single stage, continuous process with five elements in series.

=Post Treatment=

The permeate is sent to post treatment after the salt content is analyzed and controlled if too high or too low due to any unexpected errors. As stated above, the post treatment process will be done 50 meters from our plant by the Long Beach Water Reclamation Plant. The reason we aren’t performing post treatment “in-house” is due to the increased cost it would take for us to build and run the equipment for post treatment. We have the option to use the government’s water treatment facilities that already exist and our 15 million gallons per day output is a relatively small addition to the large Long Beach Water Reclamation Plant. We will assume any costs to post treatment will be pushed to the consumer in the price of purchasing the drinking water from the Long Beach Water Reclamation Plant.

Post treatment at any facility works to bring water to “taste standards”. The water that is the permeate of our process is already meeting L.A. drinking standards as shown above in Table 1 after our mass balance calculations around the reverse osmosis system (Appendix 3, Table 1). The post treatment increases the pH from 6.3 to 7.3 with the addition of base and then adds minerals like fluoride, calcium, and magnesium back in to improve taste.

==Brine Treatment and Byproduct Waste==

A natural consequence of a reverse osmosis procedure is that a lot of salt will be leftover from the membrane separation in a very concentrated form. This is called “brine” and will have to be taken care of. According to EPA requirements, the salt concentration of water entering the environment (via land or sea) cannot be higher than 40,000 ppm.8 The salt concentration of our brine is much higher (around 70,000 ppm). Thus, the brine must be removed or treated. We decided to treat the brine by diluting it as removing brine can be difficult and may require lots of land (e.g. evaporation ponds).

After the concentrate exchanges energy in the energy exchanger, it becomes a low pressure stream. It is pumped into a dilution tank and mixes with 18.2 million gallons per day of seawater from the inlet to become diluted brine which is pumped 600 meters out to sea.

The waste byproducts are the solid buildup on the filters and membranes. The filters and membranes also have a set lifetime and thus will also need to be removed. The waste is all non-hazardous and thus can be put in the municipal dump.

=Process Details and Equipment Sizing=

==Filters and UV Treatment [Pretreatment]==
The cost of our ultrafiltration system is estimated using data from the University of Tennessee at Knoxville.<ref>Qiang He and R. Bruce Robinson Membrane Filtration Systems. “Membrane Module Costs.” University of Tennessee at Knoxville. December 2014.</ref> It estimated the cost of the cost of the ultrafiltration $0.057 per 1000 gallons and the construction and capital costs are $0.39 per gallon per day. The cost of UV treatment is based off of units available from buyultraviolet.com for units that can treat about 450 gallons per minute.<ref>Atlantic Ultraviolet Corporation. “Megatron Water Disinfection System.”</ref> Due to the water having to be in relatively small pipes for treatment we do not expect the costs to scale very much, so we believe these units to be an accurate cost estimate.

==Ion Exchanger [Pretreatment] ==

The ion exchanger was sized using data from DOW. Our process will use 60 vessels with ion exchange beds of volume 5.73 m3. With a flow rate of 30 bed volumes/hour (10,312.2 m3/hr) the system of beds has a capacity of 125% of the feed. This allows for periodic regeneration of the resin, a weak acid cation resin (DOW MAC-3 Resin). This weak cation resin will remove cationic calcium and magnesium from the water to prevent scaling. The mass flow was calculated via a mass balance on the ion exchanger (Appendix 3).

==Reverse Osmosis System [Desalination]==

Each of the 2607 Reverse Osmosis SW30HRLE-440i [Seawater 30 High Rejection Low Energy 440 ft2 interlocking cap membrane] units have the same dimensions. A, the net length, is 1016 mm. B, the total length including the endcap, is 1029 mm. C, the diameter of the filter element, is 201mm. D, the diameter of the permeate pathway, is 29 mm. For a single stage, we have five of the elements in series so the overall length of a single stage is expected to be roughly 5145 mm. The membrane material is polyamide thin film composite. It will last three years.
Operating Conditions [ROSA 9.0]: 
* Temperature: 30oC 
* Pressure: 58 bar 
* Flow Rate: 52.3 Million gal per day 
* Recovery: 28.7%

[[File:Figure_4.png|center|400px|frame|Figure 4. Diagram of Dow FILMTEC SW30HRLE-440i]]
Cleaning Schedule: 
* Conservative Estimate: Once every three months (three days available for cleaning)<ref name = "Bates"/> 
* DOW FILMTEC Cleaning Conditions 
** The normalized permeate flow drops 10%
* The normalized salt passage increases 5 - 10%

==Seawater Holding Tank and Dilution Tank [Brine Treatment]==

The initial seawater holding tank works as a buffer to prevent surges and help perform shutdowns quickly and easily. We expect to run continuously for 350 days a year with the rest of the available time spread out and used for shutting down the plant for cleaning membranes, filters, and the fixing of unexpected errors. There is a big tradeoff between having a holding tank with high retention time and having a holding tank with low cost. With that in mind, we erred on the side of lowering capital cost and went with a retention time of one hour. The volume could be found by multiplying the retention time with the feed flow rate (Stream 1). We assumed that a good height to width ratio for a large water tank would be 4:1.

A mass balance was done around the dilution tank to find what amount of seawater would be necessary to dilute the brine to 45,000 ppm salinity (Appendix 3; Table 3). With this the feed flow rate into the tank could be found. It was found that it takes about 20 minutes of retention time for good dilution of the brine in seawater.8 We again assumed that a good height to width ratio for a large water tank would be 4:1.

We assumed a corrosion allowance of 4 mm wall thickness. We calculated wall thickness using vessel pressure, hoop stress, longitudinal stress, and weld factor. Each tank is made with SS316 to protect against corrosion.<ref>Towler G, Sinnott R. Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013.</ref>

<center>
'''Table 2.''' Tank sizes.
[[File:Table_2.jpg|center|600px]]
</center>

=Process Optimization=
<center> '''Table 3.'''Optimization Criteria
[[File:Table_3.jpg|center|400px]]
</center>
In order to optimize the reverse osmosis system, numerous simulations were run. Three representative cases are reported as shown above (Table 3). The objective of the optimization was to minimize the required water price to break even in 10 years. In order to achieve this goal, the balance between the capital cost from the number of the filter elements and utility cost from energy consumption was weighed. Before the energy recovery system was implemented, the general trend was that if we have more filter elements in series, overall recovery increases but recovery rate per filter decreases, and we need more filter elements. While increasing the elements in series increase the required number of filters, and thereby increases the capital cost, we require less energy to produce the same amount of water. After the energy recovery system is implemented, the correlation between the number of elements in series and the required energy becomes more complicated because high pressure concentrate is used to pressurize the feed. However, the trade-off between the utility cost and capital cost still exists. Of the three cases shown above, single stage, five elements RO system turned out to be the most optimal solution because single stage, single element system required higher utility cost and two stage system slightly did not meet the LA standard. Detailed information which was used for the optimization is available in Appendix 3 and in Appendix 6.

=Ethical, Environmental, and Societal Considerations=

There are a few key considerations we have followed to meet ethical, environmental, and societal guidelines and requirements: 
* Meet L.A. Drinking Water Standards
** Correct ion concentration
* Moved near water treatment plant
** Pipes will not corrode from low pH
* Working below pressure limits
* Water will be locally served and meet taste standards
** Via Post Treatment mineralization
* Follows EPA Environmental Salinity Requirements
* Reduced Energy Requirements
** PX Pressure Exchanger
** Optimized RO System
** Reverse Osmosis is itself Energy Saving

For further environmental considerations, see effluent summary in Appendix 4.

=Economic Analysis=

The pricing of the majority of the process was completed using the Aspen Economic Analyzer, however we had determine costs for the ultrafiltration, reverse osmosis membranes, and the ultraviolet treatment was completed separately. This led to a total ISBL of $56 MM, OSBL of $22.4 MM, Engineering Costs of $22.4 MM and a Contingency of $7.8 MM. We also chose a working Capital of $10 MM. This working capital is sufficient because we have no raw materials costs. We also had FCOP of $5.2 MM and VCOP of $3.5 MM.

We originally were planning on pricing our water at $700 per acre-foot, but after doing economic analysis we had a 10 year NPV of $-76.2 MM. After determining did a sensitivity analysis and determined that if our water was priced at $2104 per acre-foot we could have an NPV of $0 after 10 years and if we were able to get the same price of $2257 per acre-foot as the plant in Carlsbad we would have a 10 year NPV of $8.3 MM which increases to $41.1 MM at 20 years. For more detailed data about the economic analysis see Appendix 6.

=Conclusion and Recommendations=

Our mission was to create an economically viable process that will desalinate water for human consumption in the L.A. County. We have found after research and optimization that the best process to undergo desalination is a continuous single stage, five elements reverse osmosis system. Our process can meet the required L.A. drinking water standard and we have set up plans for post treatment at a local water reclamation site. While meeting the drinking water standards, we also met our ethical, environmental, and social standards.

Nevertheless, at a sale price of $700 per acre-foot, our proposed desalination plant in Los Angeles, California looks to be a very expensive project to undergo. Without any sort of government assistance to carry the heavy capital costs of such a project, we would suggest the plant not be built. However, after performing a sensitivity analysis at higher (and less conservative) price points we reached a breakeven point in 10 years with a sale price of $2104 per acre-foot of drinking water. Research has shown that the droughts striking California look only to get worse in the near future. With limited freshwater resources, there is no option other than to consider desalination. Therefore it is reasonable to expect we could get near the same price deal as local desalination plants like the Carlsbad Plant ($2257 per acre-foot) As we have met all design constraints for this project, we hope to convince policymakers to adopt our project at least $2104 per acre-foot water. If this deal is reached, we recommend that this project and design be funded.

However, there are ways to improve our design with future work. Right now all the pipes diameters are based on having a flowrate of 1.5 m/s. But in reality, we’d need to use standard piping sizes to get more accurate pricing and energy costs. Also, in reality we would be purchasing some of these vessels in bulk quantities (i.e. bulk filters) so the price could be lowered with negotiation and economies of scale. We also have pumps that are providing very little pressure for their high cost. We can remove these pumps and simply run at higher pressures earlier. We can potentially reduce early energy costs by building a “floating” desalination plant in the sea to reduce pumping distance from the well.

=Appendices=

==Appendix 1: Design Basis==
[[File:Appendix_1.png|center|Appendix_1.png]]

==Appendix 2: Process Flow Diagram==
[[File:Appendix_2.png|center]]
==Appendix 3: Mass and Energy Balance Tables==
'''Table 1'''. Mass balance on 1 stage 5 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T1.png|center|500px|]]

'''Table 2'''. Mass balance on 1 stage 1 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T2.png|center|500px|]]

'''Table 3'''. Mass balance on 2 stage 5,5 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T3.png|center|600px|]]

'''Table 4'''. 1 stage 5 elements CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T4.png|center|300px|]]

'''Table 5'''. 1 stage 1 element CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T5.png|center|300px|]]

'''Table 6'''. 2 stages 5,5 elements CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T6.png|center|300px|]]

'''Table 7'''. Energy balance on 1 stage 5 elements RO system
[[File:Appendix_3_T7.png|center|800px|]]

==Appendix 4: Effluent Summary==
[[File:Appendix_4.png|center]]

==Appendix 5: Cost of Production==
[[File:Appendix_5.png|center]]

==Appendix 6: Economic Anlaysis==
===Case I ===
Water is $700; RO with 1 stage and 5 elements
[[File:Appendix_6_C1.png|center]]

===Case II===
Water is $1400; RO with 1 stage and 5 elements
[[File:Appendix_6_C2.png|center]]

===Case III: Carlsbad ===
Water is $2257; RO with 1 stage and 5 elements
[[File:Appendix_6_C3.png|center]]

===Case IV: Breakeven (NPV = 0 at 10 years)===
Water is $2104; RO with 1 stage and 5 elements
[[File:Appendix_6_C4.png|center]]

===Case V: Breakeven (NPV = 0 at 10 years)===
Water is $2112; RO with 2 stages and 5,5 elements
[[File:Appendix_6_C5.png|center]]

===Case VI: Breakeven (NPV = 0 at 10 years)===
Water is $2138 with 1 stage and 1 element
[[File:Appendix_6_C6.png|center]]

==Appendix 7: ROSA Simulation Outputs==

'''Table 1'''. System detail of 1 stage 5 elements RO system
[[File:Appendix_7_T1.png|center|500px|]]

'''Table 2'''. System detail of 1 stage 1 elements RO system
[[File:Appendix_7_T2.png|center|500px|]]

'''Table 3'''. System detail of 2 stage 5,5 elements system
[[File:Appendix_7_T3.png|center|500px|]]

'''Table 4'''. 1 stage 5 elements RO system concentrations and pH in each stream
[[File:Appendix_7_T4.png|center|500px|]]

'''Table 5'''. 1 stage 1 element RO system concentration and pH in each stream
[[File:Appendix_7_T5.png|center|500px|]]

'''Table 6'''. 2 stage 5,5 elements RO system concentrations and pH in each stream
[[File:Appendix_7_T6.png|center|700px|]]

=References=
<references/>

Desalination - Team E

2016-03-12T03:39:13Z

Bts511: /* Appendix 7: ROSA Simulation Outputs */

Group E Corporation

Authors: Hassan Ali, Woo Soo Choe, Brett Sleyster, Jake Stolley

Instructors: Fengqi You, David Wegerer

March 11, 2016

=Executive Summary=

Outlined in this report is the project model report for the proposed desalination plant in Los Angeles, California which will provide drinking water to local citizens using reverse osmosis process. This location was chosen due to the region’s propensity for droughts in recent years. The fresh water resources in California have plummeted as water prices have risen 3-4% each year. This is only expected to continue. Other companies have worked deals with the Californian government to create desalination plants that provide a new fresh water drinking source. This suggests that there is a strong market here with potential growth.

The plant will have an output of 15 million gallons per day, which is about medium size relative to other plants. Thus the build time will be about one year. The plant will run for 30 years, with 350 days per year of use. Shutdown time will be used for cleaning and maintenance of equipment. The plant will be built on Long Beach, 500 meters from the coast nearby the Long Beach Water Reclamation Plant (LBWP). The input will withdraw 70.5 million gallons per day of seawater that will be used for reverse osmosis and dilution of brine.

Reverse osmosis was chosen as the desalination method for its low energy cost and flexibility. We designed a process that will pretreat the seawater to remove waste, colloids, bacteria, and membrane-damaging scalants. Once this seawater is treated is pumped to heavy pressures through a reverse osmosis system. After doing process optimization under various system parameter changes, we found a continuous, single stage, five element reverse osmosis process to be best. It produces a near 30% recovery. We reduce energy costs by including a pressure exchanger between our concentrate outlet and adjusted feed into the reverse osmosis system.

After performing mass and energy balances, we were able to size equipment and find our permeate composition. It turns out our process will meet L.A. standards for drinking water quality. However, post treatment will be done by the local LBWP to make the water taste better and be non-caustic to pipes. This will reduce “in-house” costs. Many environmental considerations were made, to meet our environmental standards. We meet EPA salinity requirements for our brine waste by diluting it. We also don’t use a lot of energy and thus reduce our fossil fuel consumption and CO2 generation. To be safe, we work well under max pressure limits for our equipment. We make sure we are ethical and meet all standards that our local county requires of us.

After an economic evaluation at a conservative well-water price of $700 per acre-foot, we realize that our revenue is way too low to cover our costs. Capital costs are especially large with regards to our large water tanks. We find that we reach a 10 year NPV breakpoint of $2104 per acre-foot after performing a sensitivity analysis. The Carlsbad Plant in San Diego worked out a deal for $2257 per acre-foot.

As we have met all design constraints for this project, we hope to convince policymakers to adopt our project at least $2104 per acre-foot water. If this deal is reached, we recommend that this project and design be funded. There is future work to be done to improve costs such as performing pump optimization. However, this work is minimal and won’t harm our bottom line if we reach our proposed price points.

__TOC__

=Introduction=

Due to a combination of various factors--including rising human population, climate change, increased energy consumption, and land erosion--the availability of fresh water as a resource for human consumption has been dwindling. This has had a negative impact on the world, but this impact has been especially felt in areas with natural shortages in fresh water. Desert regions, such as Saudi Arabia, have resorted to desalinating seawater at the coast and transporting up to their cities. This encompasses their entire freshwater system as very few natural water sources remain. In the United States there is a region which has shown increasing signs of following in the footsteps of Saudi Arabia and becoming fully reliant on desalination: Los Angeles, California.

=Location and Market Analysis=

In this project, Los Angeles was chosen as the location our plant, because the area has been subject to a severe drought for the last four years, as shown in Figure 1. The rapid decrease in the available water has caused permanent damages to the ground, decreased the amount of water in the reservoir from 45% to 25%, and reduced the amount of hydroelectric power produced in the state. Such deficiency in usable water has increased the water prices in Los Angeles by 6-7% annually for the last decade and is expected to rise by 3.4% annually for typical customers for the next five years. For customers who use more than 20000 gallons a year, the bill is expected to increase by 34% by 2021. This price increase is occurring because of both drought and maintenance. <ref>Pacific Institute: The California Drought. 2014.</ref>

[[File:Figure_1.jpg|right|200px|frame|Figure 1. California Drought Index.]]
As of March 2015 the Metropolitan Water District of Southern California was willing to pay as much as $700 per acre-foot of water to farmers in the Sacramento Valley, so we assume this will be the target price of the water produced by the this desalination plant.<ref>Hessel, Phil. “Dry Southern California Offers Northern Farmers Top Dollar for Water”. NBC News. March 18, 2015. </ref>

Currently, there is a plant in Carlsbad which supplies drinking water to the San Diego area. The plant can produce up to 50 million gallons per day and is the nation's largest water desalination plant. The water from this plant costs $2257 per acre-foot for the first 48000 acre-feet and about $2000 for any additional acre-foot. The county signed a contract with this plant to purchase the first 48000 acre-feet of water for the next 30 years.<ref>Carlsbad Desalination Plant. “Project Overview”. January 2016.</ref> Also, there is another desalination plant in Tampa that produces about 25 million gallons of water per day, and over 2000 other public and industrial plants with production capacity of over 300000 gallons per day. When conducting the economic analysis our plant, such numbers presented by other plants must be considered, so our project can yield positive profit.<ref>Leven, Rachel. “U.S. Desalination Industry Grows Since 2000; See as Essential to Meeting Supply Needs”. Bloomberg BNA. August 21, 2013.</ref>

The profitability of the project is naturally heavily dependent on the price of water in California. If the drought ends, water supply will increase and the water price will drop significantly, but if the drought continues, the price of water in California could rise even more rapidly and allow desalination plants to be more profitable.<ref>Southern California Public Radio. “Where is California water use decreasing?”. December 2015.</ref> It may be difficult to make this plant profitable without a guaranteed purchase or other subsidy from the local governments because the presence of pre-existing plants maintains the water price relatively low. However, external sources of profit such as guaranteed purchase or subsidy are expected because the ramifications of not having desalination plant or other sources of water may be significantly more expensive than the subsidizing a desalination plant assuming the drought continues.<ref>Weiser, Matt. “Could desalination solve California’s water problem?”. The Sacramento Bee. October 18, 2014.</ref>

=Design Basis and Technical Approach=

[[File:Figure_2.jpg|left|200px|frame|Figure 2. Vertical Well Feed]]
Considering the number of pre-existing desalination plant and the potential risk of the drought ending, our plant will be tentatively designed to produce 15 million gallons of drinkable water per day to homes and businesses in the Los Angeles county area. With a recovery rate of near 30%, and a requirement for diluting any brine, the input feed flowrate will be 70.5 million gallons per day. Following the similarly built Carlsbad plant, we expect the plant to last 30 years. The project building timeline is one year based on our relatively small plant size. We will run for 350 days a year while shutting down for cleaning and maintenance once every 3 months (based on DOW standards).<ref name = "Bates">Bates, Wayne. “Cleaning your RO”. Hydranautics. </ref> The plant will be built 500 meters from the coastline on Long Beach. A borehole will be drilled 10 meters into the ground. At this depth there should be seawater aquifers below the coast that regularly recharges itself with seawater from the ocean. A vertical well containing intake holes will be put into the borehole. The intake holes will be equipped with large size filters that can easily be replaced. At the surface, the well will be connected to a large-scale pump in a well pump house. From there, the seawater will be horizontally pumped into a seawater tank in the desalination plant. This will be considered the start of our process. The process ends after pretreatment, desalination, post treatment, and brine treatment are completed. Appendix 1 shows more details on the Design Basis.

The feed composition is assumed to be the standard seawater “average” based on the DOW Chemical ROSA 9.0 Technical Manual in Table 1. However, in reality the seawater composition at the plant site will have to be chemically tested. The final product will be based on L.A. standards for drinking water quality (Table 1). The specifications in the table are upper limits that we cannot exceed. We will aim to reach the salt level standards but any remineralization and pH treatment will be exported and paid for due to the added complexity and cost in making such a system in-house.

Due to the acidity of the desalinated water coming out of the reverse osmosis (RO) system (pH ~ 6), the pumps and pipes carrying the desalinated water for post treatment will likely corrode over time. We have decided to place the desalination plant by the Long Beach Water Reclamation Plant (about 50 meters away) to prevent pipe corrosion.

[[File:Table_1.jpg|center|200px|frame|Table 1. Concentration of the standard seawater, the output from the RO system, and LA standard.]]

=Desalination Overview and Alternatives=

We first began our research into the current landscape of the desalination industry, particularly of plants that were built in the United States. We discovered there were many processes but the main ones included: 
* Membrane Separation (Forward Osmosis and Reverse Osmosis)
* Thermal Separation (Multi Stage Flash Distillation)
* Phase Change Separation (Freeze Separation)
Each process was analyzed for their pros and cons. For our project design, we decided to go with reverse osmosis as our desalination process. This is because it is the most commonly used process in the industry, and for good reason. Unlike thermal separation, which requires huge amounts of energy to supply heat, reverse osmosis is very energy efficient. This reduces energy costs that would be spread across the plant’s lifetime.<ref name = "Republic">Republic of Mauritius NY. "Technical Aspects of Desalination Plant". Sustainable Sanitation and Water Management. </ref>

Forward osmosis is a newer technology being developed that also uses membranes to separate water and salt and may be more energy efficient overall than reverse osmosis. This is because forward osmosis doesn’t require a pump to pressurize seawater through the membranes, but rather leaves a higher salt concentrated solution (AKA draw solution) on the other side of the membrane for the seawater to be drawn into. Then a second step is required to regenerate the concentrated draw solution and produce purified water.<ref>Mallinson, Alissa. "Study Shows Forward Osmosis Desalination Not Energy Efficient." MIT News [Boston]. July 23, 2014.</ref> We decided to move away from forward osmosis because it is a process that is still being refined and developed for desalination and lacks the data, programs, and technologies that reverse osmosis users benefit from. Additionally, a recent study from MIT has claimed that due to the inherent nature of forward osmosis requiring that seawater be drawn to a higher concentrated salt solution before the water and salt is separated, it cannot possibly be more energy efficient than reverse osmosis.<ref>O'Hern, Sean C., Karnik, Rohit. Nanofiltration across Defect-Sealed Nanoporous Monolayer Graphene. American Chemical Society 15.5 (2014); 3254-260. </ref>

Other processes such as freeze separation were too restrictive in their feed and process demands and required a lot of capital. Ultimately, reverse osmosis is simply the best process that exists for desalination and has been ranked the best when performance is evaluated in terms of technical, environmental, and economic aspects. It is a process that is flexible in water quantity, operation start-up and shut-off, and site location.

Nevertheless there are disadvantages to reverse osmosis that must be considered. The feed seawater must not be too salty and too low quality. The pressure to pump the seawater is enormous. The construction time is long for large plants. The capital costs are also high when compared with other processes such as thermal separation. But all of these disadvantages are ultimately outweighed by the advantages reverse osmosis provides.<ref name = "Republic"/>

=Process Overview and Alternatives=

See Appendix 2 (Figure 1) for the Process Flow Diagram (PFD). The process itself is detailed below. The stream flow rates are in Appendix 2 (Table 1). We assumed 10 meters of pipe for every stream except the inlet and outlet streams of the entire PFD.

==Pretreatment==

After the 10 meter vertical well pumps seawater from the Long Beach coast, the feed is pumped 500 meters to a seawater tank that makes it easy to shut-down the system and reduce the large volume of flow. The level control keeps the seawater from overfilling the tank. From there, part of the seawater goes to brine treatment while the rest go to pretreatment. The purpose of pretreatment in the desalination process is to prevent equipment damage by removing particles and microorganisms from the water. A variety of pollutants -- including calcium carbonate, sulfates, silica, clay, bacteria, and other colloids and microorganisms -- can disrupt the reverse osmosis process. Media filtration uses fine grains such as sand or anthracite to filter the particles. We used a large filter to catch larger particle material (e.g. sand). However, simple filtration is not enough to prevent colloid fouling. An ultrafilter (i.e. microfiltration) is used to remove virtually all suspended matter, effectively preventing colloid fouling. The costs associated with ultrafiltration are not prohibitive in terms of the project’s budget.<ref name = "Desalination">Desalination Pre-treatment. Lenntech.</ref>

Reverse osmosis systems can also be disrupted by the growth of microbiological films of bacteria on the film and pipe walls. They are typically treated as colloids during pretreatment, and most organisms are removed from the water by ultrafiltration, but this does not remove the smallest of organic particles. DBNPA is the most effective biocide, but these aren’t often used and have to be removed from the water. Chlorination is an effective way to remove bacteria, but it can cause degradation of the membrane, and must be removed before the reverse osmosis occurs. In our process we will use UV irradiation to remove any organic material that passes through the microfilters.<ref name = "Desalination"/>

Scaling can reduce the flux across the membrane and can build up in pipe corners, causing backup. It is caused by inorganic salts -- specifically sulfates and calcium and magnesium carbonate -- present in the water. Scaling can be prevented and/or controlled with acids, antiscalants, or ion exchange softening. Antiscalants are good for removing sulfates, but not much else. Acidification will only be effective at removing CaCO3 from water but, the decreased pH may deteriorate the membrane and pipes, so it may require future research.<ref name = "Desalination"/> Ion exchange is very expensive and requires more capital to undergo, but is overall the more effective option.

The ion exchanger is a pretreatment system which acts as a method to fight membrane scaling in reverse osmosis. The ion exchanger works by using sodium carboxylate as a resin which will react with Chlorine and Magnesium ions, removing them from the feed. The sodium is substituted out and enters the feed stream, increasing the overall sodium level and TDS. The lack of magnesium and chlorine ions means that there is nothing to react with the carbonates to form scaling salts on the membranes.

==Desalination (Reverse Osmosis)==

Two pumps (P-105 & P-106) take the adjusted feed seawater up to 58 bar. However, before this pumping occurs, about 70% of the original stream gets separated and heads into a PX Pressure Exchanger. This separated stream later joins together with the pumped stream to enter the reverse osmosis system at the required pressure of 58 bar. The streams can combine together as they have equal pressure. It is the concentrate stream from the output of the reverse osmosis system that transfers its energy to the adjusted feed in the PX Pressure Exchanger.

[[File:Figure_3.jpg|left|300px|frame|Figure 3. Inside look of pressure exchanger]]
The pressure of the concentrate stream is about 57 bar and it is fed into an exchanger which works by using vessel that resembles a small pipe within a rotating wheel. The low pressure (LP) feed enters the vessel at 1 bar (see Figure 3). The feed turns and is pushed out into an outlet pipe by the high pressure (HP) concentrate stream at ~57 bar. The feed is now a HP stream via this exchanged energy. The now LP concentrate stream turns and is pushed out into an outlet pipe by the LP feed entering the vessel.

98% of the energy within the concentrate stream can be recovered via the PX Pressure Exchanger.<ref>Desalination Product and Applications Catalog 2013. “Advance Technologies for Seawater Reverse Osmosis and Brackish Water Applications”. Energy Recovery, Inc. 2012-2014.</ref >Thus, the now HP feed stream is not quite 58 bar and will need to be pumped by a small booster pump (P-107).

When designing a reverse osmosis system, we need to make decisions on what specific kind of reverse osmosis system will be implemented. When making the decision, we investigated the advantages of different types of reverse osmosis among the following options: batch vs. continuous process, number of stages, and number of elements.<ref>FILMTEC Reverse Osmosis Membranes. “Water & Process Solutions Technical Manual.” Dow Chemical Company.</ref> Our decision was made based on process optimization which is described later in this report. The end result of that optimization was to go with a single stage, continuous process with five elements in series.

=Post Treatment=

The permeate is sent to post treatment after the salt content is analyzed and controlled if too high or too low due to any unexpected errors. As stated above, the post treatment process will be done 50 meters from our plant by the Long Beach Water Reclamation Plant. The reason we aren’t performing post treatment “in-house” is due to the increased cost it would take for us to build and run the equipment for post treatment. We have the option to use the government’s water treatment facilities that already exist and our 15 million gallons per day output is a relatively small addition to the large Long Beach Water Reclamation Plant. We will assume any costs to post treatment will be pushed to the consumer in the price of purchasing the drinking water from the Long Beach Water Reclamation Plant.

Post treatment at any facility works to bring water to “taste standards”. The water that is the permeate of our process is already meeting L.A. drinking standards as shown above in Table 1 after our mass balance calculations around the reverse osmosis system (Appendix 3, Table 1). The post treatment increases the pH from 6.3 to 7.3 with the addition of base and then adds minerals like fluoride, calcium, and magnesium back in to improve taste.

==Brine Treatment and Byproduct Waste==

A natural consequence of a reverse osmosis procedure is that a lot of salt will be leftover from the membrane separation in a very concentrated form. This is called “brine” and will have to be taken care of. According to EPA requirements, the salt concentration of water entering the environment (via land or sea) cannot be higher than 40,000 ppm.8 The salt concentration of our brine is much higher (around 70,000 ppm). Thus, the brine must be removed or treated. We decided to treat the brine by diluting it as removing brine can be difficult and may require lots of land (e.g. evaporation ponds).

After the concentrate exchanges energy in the energy exchanger, it becomes a low pressure stream. It is pumped into a dilution tank and mixes with 18.2 million gallons per day of seawater from the inlet to become diluted brine which is pumped 600 meters out to sea.

The waste byproducts are the solid buildup on the filters and membranes. The filters and membranes also have a set lifetime and thus will also need to be removed. The waste is all non-hazardous and thus can be put in the municipal dump.

=Process Details and Equipment Sizing=

==Filters and UV Treatment [Pretreatment]==
The cost of our ultrafiltration system is estimated using data from the University of Tennessee at Knoxville.<ref>Qiang He and R. Bruce Robinson Membrane Filtration Systems. “Membrane Module Costs.” University of Tennessee at Knoxville. December 2014.</ref> It estimated the cost of the cost of the ultrafiltration $0.057 per 1000 gallons and the construction and capital costs are $0.39 per gallon per day. The cost of UV treatment is based off of units available from buyultraviolet.com for units that can treat about 450 gallons per minute.<ref>Atlantic Ultraviolet Corporation. “Megatron Water Disinfection System.”</ref> Due to the water having to be in relatively small pipes for treatment we do not expect the costs to scale very much, so we believe these units to be an accurate cost estimate.

==Ion Exchanger [Pretreatment] ==

The ion exchanger was sized using data from DOW. Our process will use 60 vessels with ion exchange beds of volume 5.73 m3. With a flow rate of 30 bed volumes/hour (10,312.2 m3/hr) the system of beds has a capacity of 125% of the feed. This allows for periodic regeneration of the resin, a weak acid cation resin (DOW MAC-3 Resin). This weak cation resin will remove cationic calcium and magnesium from the water to prevent scaling. The mass flow was calculated via a mass balance on the ion exchanger (Appendix 3).

==Reverse Osmosis System [Desalination]==

Each of the 2607 Reverse Osmosis SW30HRLE-440i [Seawater 30 High Rejection Low Energy 440 ft2 interlocking cap membrane] units have the same dimensions. A, the net length, is 1016 mm. B, the total length including the endcap, is 1029 mm. C, the diameter of the filter element, is 201mm. D, the diameter of the permeate pathway, is 29 mm. For a single stage, we have five of the elements in series so the overall length of a single stage is expected to be roughly 5145 mm. The membrane material is polyamide thin film composite. It will last three years.
Operating Conditions [ROSA 9.0]: 
* Temperature: 30oC 
* Pressure: 58 bar 
* Flow Rate: 52.3 Million gal per day 
* Recovery: 28.7%

[[File:Figure_4.png|center|400px|frame|Figure 4. Diagram of Dow FILMTEC SW30HRLE-440i]]
Cleaning Schedule: 
* Conservative Estimate: Once every three months (three days available for cleaning)<ref name = "Bates"/> 
* DOW FILMTEC Cleaning Conditions 
** The normalized permeate flow drops 10%
* The normalized salt passage increases 5 - 10%

==Seawater Holding Tank and Dilution Tank [Brine Treatment]==

The initial seawater holding tank works as a buffer to prevent surges and help perform shutdowns quickly and easily. We expect to run continuously for 350 days a year with the rest of the available time spread out and used for shutting down the plant for cleaning membranes, filters, and the fixing of unexpected errors. There is a big tradeoff between having a holding tank with high retention time and having a holding tank with low cost. With that in mind, we erred on the side of lowering capital cost and went with a retention time of one hour. The volume could be found by multiplying the retention time with the feed flow rate (Stream 1). We assumed that a good height to width ratio for a large water tank would be 4:1.

A mass balance was done around the dilution tank to find what amount of seawater would be necessary to dilute the brine to 45,000 ppm salinity (Appendix 3; Table 3). With this the feed flow rate into the tank could be found. It was found that it takes about 20 minutes of retention time for good dilution of the brine in seawater.8 We again assumed that a good height to width ratio for a large water tank would be 4:1.

We assumed a corrosion allowance of 4 mm wall thickness. We calculated wall thickness using vessel pressure, hoop stress, longitudinal stress, and weld factor. Each tank is made with SS316 to protect against corrosion.<ref>Towler G, Sinnott R. Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013.</ref>

<center>
'''Table 2.''' Tank sizes.
[[File:Table_2.jpg|center|600px]]
</center>

=Process Optimization=
<center> '''Table 3.'''Optimization Criteria
[[File:Table_3.jpg|center|400px]]
</center>
In order to optimize the reverse osmosis system, numerous simulations were run. Three representative cases are reported as shown above (Table 3). The objective of the optimization was to minimize the required water price to break even in 10 years. In order to achieve this goal, the balance between the capital cost from the number of the filter elements and utility cost from energy consumption was weighed. Before the energy recovery system was implemented, the general trend was that if we have more filter elements in series, overall recovery increases but recovery rate per filter decreases, and we need more filter elements. While increasing the elements in series increase the required number of filters, and thereby increases the capital cost, we require less energy to produce the same amount of water. After the energy recovery system is implemented, the correlation between the number of elements in series and the required energy becomes more complicated because high pressure concentrate is used to pressurize the feed. However, the trade-off between the utility cost and capital cost still exists. Of the three cases shown above, single stage, five elements RO system turned out to be the most optimal solution because single stage, single element system required higher utility cost and two stage system slightly did not meet the LA standard. Detailed information which was used for the optimization is available in Appendix 3 and in Appendix 6.

=Ethical, Environmental, and Societal Considerations=

There are a few key considerations we have followed to meet ethical, environmental, and societal guidelines and requirements: 
* Meet L.A. Drinking Water Standards
** Correct ion concentration
* Moved near water treatment plant
** Pipes will not corrode from low pH
* Working below pressure limits
* Water will be locally served and meet taste standards
** Via Post Treatment mineralization
* Follows EPA Environmental Salinity Requirements
* Reduced Energy Requirements
** PX Pressure Exchanger
** Optimized RO System
** Reverse Osmosis is itself Energy Saving

For further environmental considerations, see effluent summary in Appendix 4.

=Economic Analysis=

The pricing of the majority of the process was completed using the Aspen Economic Analyzer, however we had determine costs for the ultrafiltration, reverse osmosis membranes, and the ultraviolet treatment was completed separately. This led to a total ISBL of $56 MM, OSBL of $22.4 MM, Engineering Costs of $22.4 MM and a Contingency of $7.8 MM. We also chose a working Capital of $10 MM. This working capital is sufficient because we have no raw materials costs. We also had FCOP of $5.2 MM and VCOP of $3.5 MM.

We originally were planning on pricing our water at $700 per acre-foot, but after doing economic analysis we had a 10 year NPV of $-76.2 MM. After determining did a sensitivity analysis and determined that if our water was priced at $2104 per acre-foot we could have an NPV of $0 after 10 years and if we were able to get the same price of $2257 per acre-foot as the plant in Carlsbad we would have a 10 year NPV of $8.3 MM which increases to $41.1 MM at 20 years. For more detailed data about the economic analysis see Appendix 6.

=Conclusion and Recommendations=

Our mission was to create an economically viable process that will desalinate water for human consumption in the L.A. County. We have found after research and optimization that the best process to undergo desalination is a continuous single stage, five elements reverse osmosis system. Our process can meet the required L.A. drinking water standard and we have set up plans for post treatment at a local water reclamation site. While meeting the drinking water standards, we also met our ethical, environmental, and social standards.

Nevertheless, at a sale price of $700 per acre-foot, our proposed desalination plant in Los Angeles, California looks to be a very expensive project to undergo. Without any sort of government assistance to carry the heavy capital costs of such a project, we would suggest the plant not be built. However, after performing a sensitivity analysis at higher (and less conservative) price points we reached a breakeven point in 10 years with a sale price of $2104 per acre-foot of drinking water. Research has shown that the droughts striking California look only to get worse in the near future. With limited freshwater resources, there is no option other than to consider desalination. Therefore it is reasonable to expect we could get near the same price deal as local desalination plants like the Carlsbad Plant ($2257 per acre-foot) As we have met all design constraints for this project, we hope to convince policymakers to adopt our project at least $2104 per acre-foot water. If this deal is reached, we recommend that this project and design be funded.

However, there are ways to improve our design with future work. Right now all the pipes diameters are based on having a flowrate of 1.5 m/s. But in reality, we’d need to use standard piping sizes to get more accurate pricing and energy costs. Also, in reality we would be purchasing some of these vessels in bulk quantities (i.e. bulk filters) so the price could be lowered with negotiation and economies of scale. We also have pumps that are providing very little pressure for their high cost. We can remove these pumps and simply run at higher pressures earlier. We can potentially reduce early energy costs by building a “floating” desalination plant in the sea to reduce pumping distance from the well.

=Appendices=

==Appendix 1: Design Basis==
[[File:Appendix_1.png|center|Appendix_1.png]]

==Appendix 2: Process Flow Diagram==
[[File:Appendix_2.png|center]]
==Appendix 3: Mass and Energy Balance Tables==
'''Table 1'''. Mass balance on 1 stage 5 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T1.png|center|500px|]]

'''Table 2'''. Mass balance on 1 stage 1 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T2.png|center|500px|]]

'''Table 3'''. Mass balance on 2 stage 5,5 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T3.png|center|600px|]]

'''Table 4'''. 1 stage 5 elements CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T4.png|center|300px|]]

'''Table 5'''. 1 stage 1 element CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T5.png|center|300px|]]

'''Table 6'''. 2 stages 5,5 elements CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T6.png|center|300px|]]

'''Table 7'''. Energy balance on 1 stage 5 elements RO system
[[File:Appendix_3_T7.png|center|800px|]]

==Appendix 4: Effluent Summary==
[[File:Appendix_4.png|center]]

==Appendix 5: Cost of Production==
[[File:Appendix_5.png|center]]

==Appendix 6: Economic Anlaysis==
===Case I ===
Water is $700; RO with 1 stage and 5 elements
[[File:Appendix_6_C1.png|center]]

===Case II===
Water is $1400; RO with 1 stage and 5 elements
[[File:Appendix_6_C2.png|center]]

===Case III: Carlsbad ===
Water is $2257; RO with 1 stage and 5 elements
[[File:Appendix_6_C3.png|center]]

===Case IV: Breakeven (NPV = 0 at 10 years)===
Water is $2104; RO with 1 stage and 5 elements
[[File:Appendix_6_C4.png|center]]

===Case V: Breakeven (NPV = 0 at 10 years)===
Water is $2112; RO with 2 stages and 5,5 elements
[[File:Appendix_6_C5.png|center]]

===Case VI: Breakeven (NPV = 0 at 10 years)===
Water is $2138 with 1 stage and 1 element
[[File:Appendix_6_C6.png|center]]

==Appendix 7: ROSA Simulation Outputs==

'''Table 1'''. System detail of 1 stage 5 elements RO system
[[File:Appendix_7_T1.png|center|500px|]]

'''Table 2'''. System detail of 1 stage 1 elements RO system
[[File:Appendix_7_T2.png|center|500px|]]

'''Table 3'''. System detail of 2 stage 5,5 elements system
[[File:Appendix_7_T3.png|center|500px|]]

'''Table 4'''. 1 stage 5 elements RO system concentrations and pH in each stream
[[File:Appendix_7_T4.png|center|500px|]]

'''Table 5'''. 1 stage 1 element RO system concentration and pH in each stream
[[File:Appendix_7_T5.png|center|500px|]]

'''Table 6'''. 2 stage 5,5 elements RO system concentrations and pH in each stream
[[File:Appendix_7_T6.png|center|700px|]]

=References=
<references/>

Desalination - Team E

2016-03-12T03:38:25Z

Bts511: /* Appendix 7: ROSA Simulation Outputs */

Group E Corporation

Authors: Hassan Ali, Woo Soo Choe, Brett Sleyster, Jake Stolley

Instructors: Fengqi You, David Wegerer

March 11, 2016

=Executive Summary=

Outlined in this report is the project model report for the proposed desalination plant in Los Angeles, California which will provide drinking water to local citizens using reverse osmosis process. This location was chosen due to the region’s propensity for droughts in recent years. The fresh water resources in California have plummeted as water prices have risen 3-4% each year. This is only expected to continue. Other companies have worked deals with the Californian government to create desalination plants that provide a new fresh water drinking source. This suggests that there is a strong market here with potential growth.

The plant will have an output of 15 million gallons per day, which is about medium size relative to other plants. Thus the build time will be about one year. The plant will run for 30 years, with 350 days per year of use. Shutdown time will be used for cleaning and maintenance of equipment. The plant will be built on Long Beach, 500 meters from the coast nearby the Long Beach Water Reclamation Plant (LBWP). The input will withdraw 70.5 million gallons per day of seawater that will be used for reverse osmosis and dilution of brine.

Reverse osmosis was chosen as the desalination method for its low energy cost and flexibility. We designed a process that will pretreat the seawater to remove waste, colloids, bacteria, and membrane-damaging scalants. Once this seawater is treated is pumped to heavy pressures through a reverse osmosis system. After doing process optimization under various system parameter changes, we found a continuous, single stage, five element reverse osmosis process to be best. It produces a near 30% recovery. We reduce energy costs by including a pressure exchanger between our concentrate outlet and adjusted feed into the reverse osmosis system.

After performing mass and energy balances, we were able to size equipment and find our permeate composition. It turns out our process will meet L.A. standards for drinking water quality. However, post treatment will be done by the local LBWP to make the water taste better and be non-caustic to pipes. This will reduce “in-house” costs. Many environmental considerations were made, to meet our environmental standards. We meet EPA salinity requirements for our brine waste by diluting it. We also don’t use a lot of energy and thus reduce our fossil fuel consumption and CO2 generation. To be safe, we work well under max pressure limits for our equipment. We make sure we are ethical and meet all standards that our local county requires of us.

After an economic evaluation at a conservative well-water price of $700 per acre-foot, we realize that our revenue is way too low to cover our costs. Capital costs are especially large with regards to our large water tanks. We find that we reach a 10 year NPV breakpoint of $2104 per acre-foot after performing a sensitivity analysis. The Carlsbad Plant in San Diego worked out a deal for $2257 per acre-foot.

As we have met all design constraints for this project, we hope to convince policymakers to adopt our project at least $2104 per acre-foot water. If this deal is reached, we recommend that this project and design be funded. There is future work to be done to improve costs such as performing pump optimization. However, this work is minimal and won’t harm our bottom line if we reach our proposed price points.

__TOC__

=Introduction=

Due to a combination of various factors--including rising human population, climate change, increased energy consumption, and land erosion--the availability of fresh water as a resource for human consumption has been dwindling. This has had a negative impact on the world, but this impact has been especially felt in areas with natural shortages in fresh water. Desert regions, such as Saudi Arabia, have resorted to desalinating seawater at the coast and transporting up to their cities. This encompasses their entire freshwater system as very few natural water sources remain. In the United States there is a region which has shown increasing signs of following in the footsteps of Saudi Arabia and becoming fully reliant on desalination: Los Angeles, California.

=Location and Market Analysis=

In this project, Los Angeles was chosen as the location our plant, because the area has been subject to a severe drought for the last four years, as shown in Figure 1. The rapid decrease in the available water has caused permanent damages to the ground, decreased the amount of water in the reservoir from 45% to 25%, and reduced the amount of hydroelectric power produced in the state. Such deficiency in usable water has increased the water prices in Los Angeles by 6-7% annually for the last decade and is expected to rise by 3.4% annually for typical customers for the next five years. For customers who use more than 20000 gallons a year, the bill is expected to increase by 34% by 2021. This price increase is occurring because of both drought and maintenance. <ref>Pacific Institute: The California Drought. 2014.</ref>

[[File:Figure_1.jpg|right|200px|frame|Figure 1. California Drought Index.]]
As of March 2015 the Metropolitan Water District of Southern California was willing to pay as much as $700 per acre-foot of water to farmers in the Sacramento Valley, so we assume this will be the target price of the water produced by the this desalination plant.<ref>Hessel, Phil. “Dry Southern California Offers Northern Farmers Top Dollar for Water”. NBC News. March 18, 2015. </ref>

Currently, there is a plant in Carlsbad which supplies drinking water to the San Diego area. The plant can produce up to 50 million gallons per day and is the nation's largest water desalination plant. The water from this plant costs $2257 per acre-foot for the first 48000 acre-feet and about $2000 for any additional acre-foot. The county signed a contract with this plant to purchase the first 48000 acre-feet of water for the next 30 years.<ref>Carlsbad Desalination Plant. “Project Overview”. January 2016.</ref> Also, there is another desalination plant in Tampa that produces about 25 million gallons of water per day, and over 2000 other public and industrial plants with production capacity of over 300000 gallons per day. When conducting the economic analysis our plant, such numbers presented by other plants must be considered, so our project can yield positive profit.<ref>Leven, Rachel. “U.S. Desalination Industry Grows Since 2000; See as Essential to Meeting Supply Needs”. Bloomberg BNA. August 21, 2013.</ref>

The profitability of the project is naturally heavily dependent on the price of water in California. If the drought ends, water supply will increase and the water price will drop significantly, but if the drought continues, the price of water in California could rise even more rapidly and allow desalination plants to be more profitable.<ref>Southern California Public Radio. “Where is California water use decreasing?”. December 2015.</ref> It may be difficult to make this plant profitable without a guaranteed purchase or other subsidy from the local governments because the presence of pre-existing plants maintains the water price relatively low. However, external sources of profit such as guaranteed purchase or subsidy are expected because the ramifications of not having desalination plant or other sources of water may be significantly more expensive than the subsidizing a desalination plant assuming the drought continues.<ref>Weiser, Matt. “Could desalination solve California’s water problem?”. The Sacramento Bee. October 18, 2014.</ref>

=Design Basis and Technical Approach=

[[File:Figure_2.jpg|left|200px|frame|Figure 2. Vertical Well Feed]]
Considering the number of pre-existing desalination plant and the potential risk of the drought ending, our plant will be tentatively designed to produce 15 million gallons of drinkable water per day to homes and businesses in the Los Angeles county area. With a recovery rate of near 30%, and a requirement for diluting any brine, the input feed flowrate will be 70.5 million gallons per day. Following the similarly built Carlsbad plant, we expect the plant to last 30 years. The project building timeline is one year based on our relatively small plant size. We will run for 350 days a year while shutting down for cleaning and maintenance once every 3 months (based on DOW standards).<ref name = "Bates">Bates, Wayne. “Cleaning your RO”. Hydranautics. </ref> The plant will be built 500 meters from the coastline on Long Beach. A borehole will be drilled 10 meters into the ground. At this depth there should be seawater aquifers below the coast that regularly recharges itself with seawater from the ocean. A vertical well containing intake holes will be put into the borehole. The intake holes will be equipped with large size filters that can easily be replaced. At the surface, the well will be connected to a large-scale pump in a well pump house. From there, the seawater will be horizontally pumped into a seawater tank in the desalination plant. This will be considered the start of our process. The process ends after pretreatment, desalination, post treatment, and brine treatment are completed. Appendix 1 shows more details on the Design Basis.

The feed composition is assumed to be the standard seawater “average” based on the DOW Chemical ROSA 9.0 Technical Manual in Table 1. However, in reality the seawater composition at the plant site will have to be chemically tested. The final product will be based on L.A. standards for drinking water quality (Table 1). The specifications in the table are upper limits that we cannot exceed. We will aim to reach the salt level standards but any remineralization and pH treatment will be exported and paid for due to the added complexity and cost in making such a system in-house.

Due to the acidity of the desalinated water coming out of the reverse osmosis (RO) system (pH ~ 6), the pumps and pipes carrying the desalinated water for post treatment will likely corrode over time. We have decided to place the desalination plant by the Long Beach Water Reclamation Plant (about 50 meters away) to prevent pipe corrosion.

[[File:Table_1.jpg|center|200px|frame|Table 1. Concentration of the standard seawater, the output from the RO system, and LA standard.]]

=Desalination Overview and Alternatives=

We first began our research into the current landscape of the desalination industry, particularly of plants that were built in the United States. We discovered there were many processes but the main ones included: 
* Membrane Separation (Forward Osmosis and Reverse Osmosis)
* Thermal Separation (Multi Stage Flash Distillation)
* Phase Change Separation (Freeze Separation)
Each process was analyzed for their pros and cons. For our project design, we decided to go with reverse osmosis as our desalination process. This is because it is the most commonly used process in the industry, and for good reason. Unlike thermal separation, which requires huge amounts of energy to supply heat, reverse osmosis is very energy efficient. This reduces energy costs that would be spread across the plant’s lifetime.<ref name = "Republic">Republic of Mauritius NY. "Technical Aspects of Desalination Plant". Sustainable Sanitation and Water Management. </ref>

Forward osmosis is a newer technology being developed that also uses membranes to separate water and salt and may be more energy efficient overall than reverse osmosis. This is because forward osmosis doesn’t require a pump to pressurize seawater through the membranes, but rather leaves a higher salt concentrated solution (AKA draw solution) on the other side of the membrane for the seawater to be drawn into. Then a second step is required to regenerate the concentrated draw solution and produce purified water.<ref>Mallinson, Alissa. "Study Shows Forward Osmosis Desalination Not Energy Efficient." MIT News [Boston]. July 23, 2014.</ref> We decided to move away from forward osmosis because it is a process that is still being refined and developed for desalination and lacks the data, programs, and technologies that reverse osmosis users benefit from. Additionally, a recent study from MIT has claimed that due to the inherent nature of forward osmosis requiring that seawater be drawn to a higher concentrated salt solution before the water and salt is separated, it cannot possibly be more energy efficient than reverse osmosis.<ref>O'Hern, Sean C., Karnik, Rohit. Nanofiltration across Defect-Sealed Nanoporous Monolayer Graphene. American Chemical Society 15.5 (2014); 3254-260. </ref>

Other processes such as freeze separation were too restrictive in their feed and process demands and required a lot of capital. Ultimately, reverse osmosis is simply the best process that exists for desalination and has been ranked the best when performance is evaluated in terms of technical, environmental, and economic aspects. It is a process that is flexible in water quantity, operation start-up and shut-off, and site location.

Nevertheless there are disadvantages to reverse osmosis that must be considered. The feed seawater must not be too salty and too low quality. The pressure to pump the seawater is enormous. The construction time is long for large plants. The capital costs are also high when compared with other processes such as thermal separation. But all of these disadvantages are ultimately outweighed by the advantages reverse osmosis provides.<ref name = "Republic"/>

=Process Overview and Alternatives=

See Appendix 2 (Figure 1) for the Process Flow Diagram (PFD). The process itself is detailed below. The stream flow rates are in Appendix 2 (Table 1). We assumed 10 meters of pipe for every stream except the inlet and outlet streams of the entire PFD.

==Pretreatment==

After the 10 meter vertical well pumps seawater from the Long Beach coast, the feed is pumped 500 meters to a seawater tank that makes it easy to shut-down the system and reduce the large volume of flow. The level control keeps the seawater from overfilling the tank. From there, part of the seawater goes to brine treatment while the rest go to pretreatment. The purpose of pretreatment in the desalination process is to prevent equipment damage by removing particles and microorganisms from the water. A variety of pollutants -- including calcium carbonate, sulfates, silica, clay, bacteria, and other colloids and microorganisms -- can disrupt the reverse osmosis process. Media filtration uses fine grains such as sand or anthracite to filter the particles. We used a large filter to catch larger particle material (e.g. sand). However, simple filtration is not enough to prevent colloid fouling. An ultrafilter (i.e. microfiltration) is used to remove virtually all suspended matter, effectively preventing colloid fouling. The costs associated with ultrafiltration are not prohibitive in terms of the project’s budget.<ref name = "Desalination">Desalination Pre-treatment. Lenntech.</ref>

Reverse osmosis systems can also be disrupted by the growth of microbiological films of bacteria on the film and pipe walls. They are typically treated as colloids during pretreatment, and most organisms are removed from the water by ultrafiltration, but this does not remove the smallest of organic particles. DBNPA is the most effective biocide, but these aren’t often used and have to be removed from the water. Chlorination is an effective way to remove bacteria, but it can cause degradation of the membrane, and must be removed before the reverse osmosis occurs. In our process we will use UV irradiation to remove any organic material that passes through the microfilters.<ref name = "Desalination"/>

Scaling can reduce the flux across the membrane and can build up in pipe corners, causing backup. It is caused by inorganic salts -- specifically sulfates and calcium and magnesium carbonate -- present in the water. Scaling can be prevented and/or controlled with acids, antiscalants, or ion exchange softening. Antiscalants are good for removing sulfates, but not much else. Acidification will only be effective at removing CaCO3 from water but, the decreased pH may deteriorate the membrane and pipes, so it may require future research.<ref name = "Desalination"/> Ion exchange is very expensive and requires more capital to undergo, but is overall the more effective option.

The ion exchanger is a pretreatment system which acts as a method to fight membrane scaling in reverse osmosis. The ion exchanger works by using sodium carboxylate as a resin which will react with Chlorine and Magnesium ions, removing them from the feed. The sodium is substituted out and enters the feed stream, increasing the overall sodium level and TDS. The lack of magnesium and chlorine ions means that there is nothing to react with the carbonates to form scaling salts on the membranes.

==Desalination (Reverse Osmosis)==

Two pumps (P-105 & P-106) take the adjusted feed seawater up to 58 bar. However, before this pumping occurs, about 70% of the original stream gets separated and heads into a PX Pressure Exchanger. This separated stream later joins together with the pumped stream to enter the reverse osmosis system at the required pressure of 58 bar. The streams can combine together as they have equal pressure. It is the concentrate stream from the output of the reverse osmosis system that transfers its energy to the adjusted feed in the PX Pressure Exchanger.

[[File:Figure_3.jpg|left|300px|frame|Figure 3. Inside look of pressure exchanger]]
The pressure of the concentrate stream is about 57 bar and it is fed into an exchanger which works by using vessel that resembles a small pipe within a rotating wheel. The low pressure (LP) feed enters the vessel at 1 bar (see Figure 3). The feed turns and is pushed out into an outlet pipe by the high pressure (HP) concentrate stream at ~57 bar. The feed is now a HP stream via this exchanged energy. The now LP concentrate stream turns and is pushed out into an outlet pipe by the LP feed entering the vessel.

98% of the energy within the concentrate stream can be recovered via the PX Pressure Exchanger.<ref>Desalination Product and Applications Catalog 2013. “Advance Technologies for Seawater Reverse Osmosis and Brackish Water Applications”. Energy Recovery, Inc. 2012-2014.</ref >Thus, the now HP feed stream is not quite 58 bar and will need to be pumped by a small booster pump (P-107).

When designing a reverse osmosis system, we need to make decisions on what specific kind of reverse osmosis system will be implemented. When making the decision, we investigated the advantages of different types of reverse osmosis among the following options: batch vs. continuous process, number of stages, and number of elements.<ref>FILMTEC Reverse Osmosis Membranes. “Water & Process Solutions Technical Manual.” Dow Chemical Company.</ref> Our decision was made based on process optimization which is described later in this report. The end result of that optimization was to go with a single stage, continuous process with five elements in series.

=Post Treatment=

The permeate is sent to post treatment after the salt content is analyzed and controlled if too high or too low due to any unexpected errors. As stated above, the post treatment process will be done 50 meters from our plant by the Long Beach Water Reclamation Plant. The reason we aren’t performing post treatment “in-house” is due to the increased cost it would take for us to build and run the equipment for post treatment. We have the option to use the government’s water treatment facilities that already exist and our 15 million gallons per day output is a relatively small addition to the large Long Beach Water Reclamation Plant. We will assume any costs to post treatment will be pushed to the consumer in the price of purchasing the drinking water from the Long Beach Water Reclamation Plant.

Post treatment at any facility works to bring water to “taste standards”. The water that is the permeate of our process is already meeting L.A. drinking standards as shown above in Table 1 after our mass balance calculations around the reverse osmosis system (Appendix 3, Table 1). The post treatment increases the pH from 6.3 to 7.3 with the addition of base and then adds minerals like fluoride, calcium, and magnesium back in to improve taste.

==Brine Treatment and Byproduct Waste==

A natural consequence of a reverse osmosis procedure is that a lot of salt will be leftover from the membrane separation in a very concentrated form. This is called “brine” and will have to be taken care of. According to EPA requirements, the salt concentration of water entering the environment (via land or sea) cannot be higher than 40,000 ppm.8 The salt concentration of our brine is much higher (around 70,000 ppm). Thus, the brine must be removed or treated. We decided to treat the brine by diluting it as removing brine can be difficult and may require lots of land (e.g. evaporation ponds).

After the concentrate exchanges energy in the energy exchanger, it becomes a low pressure stream. It is pumped into a dilution tank and mixes with 18.2 million gallons per day of seawater from the inlet to become diluted brine which is pumped 600 meters out to sea.

The waste byproducts are the solid buildup on the filters and membranes. The filters and membranes also have a set lifetime and thus will also need to be removed. The waste is all non-hazardous and thus can be put in the municipal dump.

=Process Details and Equipment Sizing=

==Filters and UV Treatment [Pretreatment]==
The cost of our ultrafiltration system is estimated using data from the University of Tennessee at Knoxville.<ref>Qiang He and R. Bruce Robinson Membrane Filtration Systems. “Membrane Module Costs.” University of Tennessee at Knoxville. December 2014.</ref> It estimated the cost of the cost of the ultrafiltration $0.057 per 1000 gallons and the construction and capital costs are $0.39 per gallon per day. The cost of UV treatment is based off of units available from buyultraviolet.com for units that can treat about 450 gallons per minute.<ref>Atlantic Ultraviolet Corporation. “Megatron Water Disinfection System.”</ref> Due to the water having to be in relatively small pipes for treatment we do not expect the costs to scale very much, so we believe these units to be an accurate cost estimate.

==Ion Exchanger [Pretreatment] ==

The ion exchanger was sized using data from DOW. Our process will use 60 vessels with ion exchange beds of volume 5.73 m3. With a flow rate of 30 bed volumes/hour (10,312.2 m3/hr) the system of beds has a capacity of 125% of the feed. This allows for periodic regeneration of the resin, a weak acid cation resin (DOW MAC-3 Resin). This weak cation resin will remove cationic calcium and magnesium from the water to prevent scaling. The mass flow was calculated via a mass balance on the ion exchanger (Appendix 3).

==Reverse Osmosis System [Desalination]==

Each of the 2607 Reverse Osmosis SW30HRLE-440i [Seawater 30 High Rejection Low Energy 440 ft2 interlocking cap membrane] units have the same dimensions. A, the net length, is 1016 mm. B, the total length including the endcap, is 1029 mm. C, the diameter of the filter element, is 201mm. D, the diameter of the permeate pathway, is 29 mm. For a single stage, we have five of the elements in series so the overall length of a single stage is expected to be roughly 5145 mm. The membrane material is polyamide thin film composite. It will last three years.
Operating Conditions [ROSA 9.0]: 
* Temperature: 30oC 
* Pressure: 58 bar 
* Flow Rate: 52.3 Million gal per day 
* Recovery: 28.7%

[[File:Figure_4.png|center|400px|frame|Figure 4. Diagram of Dow FILMTEC SW30HRLE-440i]]
Cleaning Schedule: 
* Conservative Estimate: Once every three months (three days available for cleaning)<ref name = "Bates"/> 
* DOW FILMTEC Cleaning Conditions 
** The normalized permeate flow drops 10%
* The normalized salt passage increases 5 - 10%

==Seawater Holding Tank and Dilution Tank [Brine Treatment]==

The initial seawater holding tank works as a buffer to prevent surges and help perform shutdowns quickly and easily. We expect to run continuously for 350 days a year with the rest of the available time spread out and used for shutting down the plant for cleaning membranes, filters, and the fixing of unexpected errors. There is a big tradeoff between having a holding tank with high retention time and having a holding tank with low cost. With that in mind, we erred on the side of lowering capital cost and went with a retention time of one hour. The volume could be found by multiplying the retention time with the feed flow rate (Stream 1). We assumed that a good height to width ratio for a large water tank would be 4:1.

A mass balance was done around the dilution tank to find what amount of seawater would be necessary to dilute the brine to 45,000 ppm salinity (Appendix 3; Table 3). With this the feed flow rate into the tank could be found. It was found that it takes about 20 minutes of retention time for good dilution of the brine in seawater.8 We again assumed that a good height to width ratio for a large water tank would be 4:1.

We assumed a corrosion allowance of 4 mm wall thickness. We calculated wall thickness using vessel pressure, hoop stress, longitudinal stress, and weld factor. Each tank is made with SS316 to protect against corrosion.<ref>Towler G, Sinnott R. Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013.</ref>

<center>
'''Table 2.''' Tank sizes.
[[File:Table_2.jpg|center|600px]]
</center>

=Process Optimization=
<center> '''Table 3.'''Optimization Criteria
[[File:Table_3.jpg|center|400px]]
</center>
In order to optimize the reverse osmosis system, numerous simulations were run. Three representative cases are reported as shown above (Table 3). The objective of the optimization was to minimize the required water price to break even in 10 years. In order to achieve this goal, the balance between the capital cost from the number of the filter elements and utility cost from energy consumption was weighed. Before the energy recovery system was implemented, the general trend was that if we have more filter elements in series, overall recovery increases but recovery rate per filter decreases, and we need more filter elements. While increasing the elements in series increase the required number of filters, and thereby increases the capital cost, we require less energy to produce the same amount of water. After the energy recovery system is implemented, the correlation between the number of elements in series and the required energy becomes more complicated because high pressure concentrate is used to pressurize the feed. However, the trade-off between the utility cost and capital cost still exists. Of the three cases shown above, single stage, five elements RO system turned out to be the most optimal solution because single stage, single element system required higher utility cost and two stage system slightly did not meet the LA standard. Detailed information which was used for the optimization is available in Appendix 3 and in Appendix 6.

=Ethical, Environmental, and Societal Considerations=

There are a few key considerations we have followed to meet ethical, environmental, and societal guidelines and requirements: 
* Meet L.A. Drinking Water Standards
** Correct ion concentration
* Moved near water treatment plant
** Pipes will not corrode from low pH
* Working below pressure limits
* Water will be locally served and meet taste standards
** Via Post Treatment mineralization
* Follows EPA Environmental Salinity Requirements
* Reduced Energy Requirements
** PX Pressure Exchanger
** Optimized RO System
** Reverse Osmosis is itself Energy Saving

For further environmental considerations, see effluent summary in Appendix 4.

=Economic Analysis=

The pricing of the majority of the process was completed using the Aspen Economic Analyzer, however we had determine costs for the ultrafiltration, reverse osmosis membranes, and the ultraviolet treatment was completed separately. This led to a total ISBL of $56 MM, OSBL of $22.4 MM, Engineering Costs of $22.4 MM and a Contingency of $7.8 MM. We also chose a working Capital of $10 MM. This working capital is sufficient because we have no raw materials costs. We also had FCOP of $5.2 MM and VCOP of $3.5 MM.

We originally were planning on pricing our water at $700 per acre-foot, but after doing economic analysis we had a 10 year NPV of $-76.2 MM. After determining did a sensitivity analysis and determined that if our water was priced at $2104 per acre-foot we could have an NPV of $0 after 10 years and if we were able to get the same price of $2257 per acre-foot as the plant in Carlsbad we would have a 10 year NPV of $8.3 MM which increases to $41.1 MM at 20 years. For more detailed data about the economic analysis see Appendix 6.

=Conclusion and Recommendations=

Our mission was to create an economically viable process that will desalinate water for human consumption in the L.A. County. We have found after research and optimization that the best process to undergo desalination is a continuous single stage, five elements reverse osmosis system. Our process can meet the required L.A. drinking water standard and we have set up plans for post treatment at a local water reclamation site. While meeting the drinking water standards, we also met our ethical, environmental, and social standards.

Nevertheless, at a sale price of $700 per acre-foot, our proposed desalination plant in Los Angeles, California looks to be a very expensive project to undergo. Without any sort of government assistance to carry the heavy capital costs of such a project, we would suggest the plant not be built. However, after performing a sensitivity analysis at higher (and less conservative) price points we reached a breakeven point in 10 years with a sale price of $2104 per acre-foot of drinking water. Research has shown that the droughts striking California look only to get worse in the near future. With limited freshwater resources, there is no option other than to consider desalination. Therefore it is reasonable to expect we could get near the same price deal as local desalination plants like the Carlsbad Plant ($2257 per acre-foot) As we have met all design constraints for this project, we hope to convince policymakers to adopt our project at least $2104 per acre-foot water. If this deal is reached, we recommend that this project and design be funded.

However, there are ways to improve our design with future work. Right now all the pipes diameters are based on having a flowrate of 1.5 m/s. But in reality, we’d need to use standard piping sizes to get more accurate pricing and energy costs. Also, in reality we would be purchasing some of these vessels in bulk quantities (i.e. bulk filters) so the price could be lowered with negotiation and economies of scale. We also have pumps that are providing very little pressure for their high cost. We can remove these pumps and simply run at higher pressures earlier. We can potentially reduce early energy costs by building a “floating” desalination plant in the sea to reduce pumping distance from the well.

=Appendices=

==Appendix 1: Design Basis==
[[File:Appendix_1.png|center|Appendix_1.png]]

==Appendix 2: Process Flow Diagram==
[[File:Appendix_2.png|center]]
==Appendix 3: Mass and Energy Balance Tables==
'''Table 1'''. Mass balance on 1 stage 5 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T1.png|center|500px|]]

'''Table 2'''. Mass balance on 1 stage 1 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T2.png|center|500px|]]

'''Table 3'''. Mass balance on 2 stage 5,5 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T3.png|center|600px|]]

'''Table 4'''. 1 stage 5 elements CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T4.png|center|300px|]]

'''Table 5'''. 1 stage 1 element CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T5.png|center|300px|]]

'''Table 6'''. 2 stages 5,5 elements CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T6.png|center|300px|]]

'''Table 7'''. Energy balance on 1 stage 5 elements RO system
[[File:Appendix_3_T7.png|center|800px|]]

==Appendix 4: Effluent Summary==
[[File:Appendix_4.png|center]]

==Appendix 5: Cost of Production==
[[File:Appendix_5.png|center]]

==Appendix 6: Economic Anlaysis==
===Case I ===
Water is $700; RO with 1 stage and 5 elements
[[File:Appendix_6_C1.png|center]]

===Case II===
Water is $1400; RO with 1 stage and 5 elements
[[File:Appendix_6_C2.png|center]]

===Case III: Carlsbad ===
Water is $2257; RO with 1 stage and 5 elements
[[File:Appendix_6_C3.png|center]]

===Case IV: Breakeven (NPV = 0 at 10 years)===
Water is $2104; RO with 1 stage and 5 elements
[[File:Appendix_6_C4.png|center]]

===Case V: Breakeven (NPV = 0 at 10 years)===
Water is $2112; RO with 2 stages and 5,5 elements
[[File:Appendix_6_C5.png|center]]

===Case VI: Breakeven (NPV = 0 at 10 years)===
Water is $2138 with 1 stage and 1 element
[[File:Appendix_6_C6.png|center]]

==Appendix 7: ROSA Simulation Outputs==

'''Table 1'''. System detail of 1 stage 5 elements RO system
[[File:Appendix_7_T1.png|center|500px|]]

'''Table 2'''. System detail of 1 stage 1 elements RO system
[[File:Appendix_7_T2.png|center|500px|]]

'''Table 3'''. System detail of 2 stage 5,5 elements system
[[File:Appendix_7_T3.png|center|500px|]]

'''Table 4'''. 1 stage 5 elements RO system concentrations and pH in each stream
[[File:Appendix_7_T4.png|center|500px|]]

'''Table 5'''. 1 stage 1 element RO system concentration and pH in each stream
[[File:Appendix_7_T5.png|center|500px|]]

'''Table 6'''. 2 stage 5,5 elements RO system concentrations and pH in each stream
[[File:Appendix_7_T6.png|center|500px|]]

=References=
<references/>

Desalination - Team E

2016-03-12T03:37:45Z

Bts511: /* Appendix 7: ROSA Simulation Outputs */

Group E Corporation

Authors: Hassan Ali, Woo Soo Choe, Brett Sleyster, Jake Stolley

Instructors: Fengqi You, David Wegerer

March 11, 2016

=Executive Summary=

Outlined in this report is the project model report for the proposed desalination plant in Los Angeles, California which will provide drinking water to local citizens using reverse osmosis process. This location was chosen due to the region’s propensity for droughts in recent years. The fresh water resources in California have plummeted as water prices have risen 3-4% each year. This is only expected to continue. Other companies have worked deals with the Californian government to create desalination plants that provide a new fresh water drinking source. This suggests that there is a strong market here with potential growth.

The plant will have an output of 15 million gallons per day, which is about medium size relative to other plants. Thus the build time will be about one year. The plant will run for 30 years, with 350 days per year of use. Shutdown time will be used for cleaning and maintenance of equipment. The plant will be built on Long Beach, 500 meters from the coast nearby the Long Beach Water Reclamation Plant (LBWP). The input will withdraw 70.5 million gallons per day of seawater that will be used for reverse osmosis and dilution of brine.

Reverse osmosis was chosen as the desalination method for its low energy cost and flexibility. We designed a process that will pretreat the seawater to remove waste, colloids, bacteria, and membrane-damaging scalants. Once this seawater is treated is pumped to heavy pressures through a reverse osmosis system. After doing process optimization under various system parameter changes, we found a continuous, single stage, five element reverse osmosis process to be best. It produces a near 30% recovery. We reduce energy costs by including a pressure exchanger between our concentrate outlet and adjusted feed into the reverse osmosis system.

After performing mass and energy balances, we were able to size equipment and find our permeate composition. It turns out our process will meet L.A. standards for drinking water quality. However, post treatment will be done by the local LBWP to make the water taste better and be non-caustic to pipes. This will reduce “in-house” costs. Many environmental considerations were made, to meet our environmental standards. We meet EPA salinity requirements for our brine waste by diluting it. We also don’t use a lot of energy and thus reduce our fossil fuel consumption and CO2 generation. To be safe, we work well under max pressure limits for our equipment. We make sure we are ethical and meet all standards that our local county requires of us.

After an economic evaluation at a conservative well-water price of $700 per acre-foot, we realize that our revenue is way too low to cover our costs. Capital costs are especially large with regards to our large water tanks. We find that we reach a 10 year NPV breakpoint of $2104 per acre-foot after performing a sensitivity analysis. The Carlsbad Plant in San Diego worked out a deal for $2257 per acre-foot.

As we have met all design constraints for this project, we hope to convince policymakers to adopt our project at least $2104 per acre-foot water. If this deal is reached, we recommend that this project and design be funded. There is future work to be done to improve costs such as performing pump optimization. However, this work is minimal and won’t harm our bottom line if we reach our proposed price points.

__TOC__

=Introduction=

Due to a combination of various factors--including rising human population, climate change, increased energy consumption, and land erosion--the availability of fresh water as a resource for human consumption has been dwindling. This has had a negative impact on the world, but this impact has been especially felt in areas with natural shortages in fresh water. Desert regions, such as Saudi Arabia, have resorted to desalinating seawater at the coast and transporting up to their cities. This encompasses their entire freshwater system as very few natural water sources remain. In the United States there is a region which has shown increasing signs of following in the footsteps of Saudi Arabia and becoming fully reliant on desalination: Los Angeles, California.

=Location and Market Analysis=

In this project, Los Angeles was chosen as the location our plant, because the area has been subject to a severe drought for the last four years, as shown in Figure 1. The rapid decrease in the available water has caused permanent damages to the ground, decreased the amount of water in the reservoir from 45% to 25%, and reduced the amount of hydroelectric power produced in the state. Such deficiency in usable water has increased the water prices in Los Angeles by 6-7% annually for the last decade and is expected to rise by 3.4% annually for typical customers for the next five years. For customers who use more than 20000 gallons a year, the bill is expected to increase by 34% by 2021. This price increase is occurring because of both drought and maintenance. <ref>Pacific Institute: The California Drought. 2014.</ref>

[[File:Figure_1.jpg|right|200px|frame|Figure 1. California Drought Index.]]
As of March 2015 the Metropolitan Water District of Southern California was willing to pay as much as $700 per acre-foot of water to farmers in the Sacramento Valley, so we assume this will be the target price of the water produced by the this desalination plant.<ref>Hessel, Phil. “Dry Southern California Offers Northern Farmers Top Dollar for Water”. NBC News. March 18, 2015. </ref>

Currently, there is a plant in Carlsbad which supplies drinking water to the San Diego area. The plant can produce up to 50 million gallons per day and is the nation's largest water desalination plant. The water from this plant costs $2257 per acre-foot for the first 48000 acre-feet and about $2000 for any additional acre-foot. The county signed a contract with this plant to purchase the first 48000 acre-feet of water for the next 30 years.<ref>Carlsbad Desalination Plant. “Project Overview”. January 2016.</ref> Also, there is another desalination plant in Tampa that produces about 25 million gallons of water per day, and over 2000 other public and industrial plants with production capacity of over 300000 gallons per day. When conducting the economic analysis our plant, such numbers presented by other plants must be considered, so our project can yield positive profit.<ref>Leven, Rachel. “U.S. Desalination Industry Grows Since 2000; See as Essential to Meeting Supply Needs”. Bloomberg BNA. August 21, 2013.</ref>

The profitability of the project is naturally heavily dependent on the price of water in California. If the drought ends, water supply will increase and the water price will drop significantly, but if the drought continues, the price of water in California could rise even more rapidly and allow desalination plants to be more profitable.<ref>Southern California Public Radio. “Where is California water use decreasing?”. December 2015.</ref> It may be difficult to make this plant profitable without a guaranteed purchase or other subsidy from the local governments because the presence of pre-existing plants maintains the water price relatively low. However, external sources of profit such as guaranteed purchase or subsidy are expected because the ramifications of not having desalination plant or other sources of water may be significantly more expensive than the subsidizing a desalination plant assuming the drought continues.<ref>Weiser, Matt. “Could desalination solve California’s water problem?”. The Sacramento Bee. October 18, 2014.</ref>

=Design Basis and Technical Approach=

[[File:Figure_2.jpg|left|200px|frame|Figure 2. Vertical Well Feed]]
Considering the number of pre-existing desalination plant and the potential risk of the drought ending, our plant will be tentatively designed to produce 15 million gallons of drinkable water per day to homes and businesses in the Los Angeles county area. With a recovery rate of near 30%, and a requirement for diluting any brine, the input feed flowrate will be 70.5 million gallons per day. Following the similarly built Carlsbad plant, we expect the plant to last 30 years. The project building timeline is one year based on our relatively small plant size. We will run for 350 days a year while shutting down for cleaning and maintenance once every 3 months (based on DOW standards).<ref name = "Bates">Bates, Wayne. “Cleaning your RO”. Hydranautics. </ref> The plant will be built 500 meters from the coastline on Long Beach. A borehole will be drilled 10 meters into the ground. At this depth there should be seawater aquifers below the coast that regularly recharges itself with seawater from the ocean. A vertical well containing intake holes will be put into the borehole. The intake holes will be equipped with large size filters that can easily be replaced. At the surface, the well will be connected to a large-scale pump in a well pump house. From there, the seawater will be horizontally pumped into a seawater tank in the desalination plant. This will be considered the start of our process. The process ends after pretreatment, desalination, post treatment, and brine treatment are completed. Appendix 1 shows more details on the Design Basis.

The feed composition is assumed to be the standard seawater “average” based on the DOW Chemical ROSA 9.0 Technical Manual in Table 1. However, in reality the seawater composition at the plant site will have to be chemically tested. The final product will be based on L.A. standards for drinking water quality (Table 1). The specifications in the table are upper limits that we cannot exceed. We will aim to reach the salt level standards but any remineralization and pH treatment will be exported and paid for due to the added complexity and cost in making such a system in-house.

Due to the acidity of the desalinated water coming out of the reverse osmosis (RO) system (pH ~ 6), the pumps and pipes carrying the desalinated water for post treatment will likely corrode over time. We have decided to place the desalination plant by the Long Beach Water Reclamation Plant (about 50 meters away) to prevent pipe corrosion.

[[File:Table_1.jpg|center|200px|frame|Table 1. Concentration of the standard seawater, the output from the RO system, and LA standard.]]

=Desalination Overview and Alternatives=

We first began our research into the current landscape of the desalination industry, particularly of plants that were built in the United States. We discovered there were many processes but the main ones included: 
* Membrane Separation (Forward Osmosis and Reverse Osmosis)
* Thermal Separation (Multi Stage Flash Distillation)
* Phase Change Separation (Freeze Separation)
Each process was analyzed for their pros and cons. For our project design, we decided to go with reverse osmosis as our desalination process. This is because it is the most commonly used process in the industry, and for good reason. Unlike thermal separation, which requires huge amounts of energy to supply heat, reverse osmosis is very energy efficient. This reduces energy costs that would be spread across the plant’s lifetime.<ref name = "Republic">Republic of Mauritius NY. "Technical Aspects of Desalination Plant". Sustainable Sanitation and Water Management. </ref>

Forward osmosis is a newer technology being developed that also uses membranes to separate water and salt and may be more energy efficient overall than reverse osmosis. This is because forward osmosis doesn’t require a pump to pressurize seawater through the membranes, but rather leaves a higher salt concentrated solution (AKA draw solution) on the other side of the membrane for the seawater to be drawn into. Then a second step is required to regenerate the concentrated draw solution and produce purified water.<ref>Mallinson, Alissa. "Study Shows Forward Osmosis Desalination Not Energy Efficient." MIT News [Boston]. July 23, 2014.</ref> We decided to move away from forward osmosis because it is a process that is still being refined and developed for desalination and lacks the data, programs, and technologies that reverse osmosis users benefit from. Additionally, a recent study from MIT has claimed that due to the inherent nature of forward osmosis requiring that seawater be drawn to a higher concentrated salt solution before the water and salt is separated, it cannot possibly be more energy efficient than reverse osmosis.<ref>O'Hern, Sean C., Karnik, Rohit. Nanofiltration across Defect-Sealed Nanoporous Monolayer Graphene. American Chemical Society 15.5 (2014); 3254-260. </ref>

Other processes such as freeze separation were too restrictive in their feed and process demands and required a lot of capital. Ultimately, reverse osmosis is simply the best process that exists for desalination and has been ranked the best when performance is evaluated in terms of technical, environmental, and economic aspects. It is a process that is flexible in water quantity, operation start-up and shut-off, and site location.

Nevertheless there are disadvantages to reverse osmosis that must be considered. The feed seawater must not be too salty and too low quality. The pressure to pump the seawater is enormous. The construction time is long for large plants. The capital costs are also high when compared with other processes such as thermal separation. But all of these disadvantages are ultimately outweighed by the advantages reverse osmosis provides.<ref name = "Republic"/>

=Process Overview and Alternatives=

See Appendix 2 (Figure 1) for the Process Flow Diagram (PFD). The process itself is detailed below. The stream flow rates are in Appendix 2 (Table 1). We assumed 10 meters of pipe for every stream except the inlet and outlet streams of the entire PFD.

==Pretreatment==

After the 10 meter vertical well pumps seawater from the Long Beach coast, the feed is pumped 500 meters to a seawater tank that makes it easy to shut-down the system and reduce the large volume of flow. The level control keeps the seawater from overfilling the tank. From there, part of the seawater goes to brine treatment while the rest go to pretreatment. The purpose of pretreatment in the desalination process is to prevent equipment damage by removing particles and microorganisms from the water. A variety of pollutants -- including calcium carbonate, sulfates, silica, clay, bacteria, and other colloids and microorganisms -- can disrupt the reverse osmosis process. Media filtration uses fine grains such as sand or anthracite to filter the particles. We used a large filter to catch larger particle material (e.g. sand). However, simple filtration is not enough to prevent colloid fouling. An ultrafilter (i.e. microfiltration) is used to remove virtually all suspended matter, effectively preventing colloid fouling. The costs associated with ultrafiltration are not prohibitive in terms of the project’s budget.<ref name = "Desalination">Desalination Pre-treatment. Lenntech.</ref>

Reverse osmosis systems can also be disrupted by the growth of microbiological films of bacteria on the film and pipe walls. They are typically treated as colloids during pretreatment, and most organisms are removed from the water by ultrafiltration, but this does not remove the smallest of organic particles. DBNPA is the most effective biocide, but these aren’t often used and have to be removed from the water. Chlorination is an effective way to remove bacteria, but it can cause degradation of the membrane, and must be removed before the reverse osmosis occurs. In our process we will use UV irradiation to remove any organic material that passes through the microfilters.<ref name = "Desalination"/>

Scaling can reduce the flux across the membrane and can build up in pipe corners, causing backup. It is caused by inorganic salts -- specifically sulfates and calcium and magnesium carbonate -- present in the water. Scaling can be prevented and/or controlled with acids, antiscalants, or ion exchange softening. Antiscalants are good for removing sulfates, but not much else. Acidification will only be effective at removing CaCO3 from water but, the decreased pH may deteriorate the membrane and pipes, so it may require future research.<ref name = "Desalination"/> Ion exchange is very expensive and requires more capital to undergo, but is overall the more effective option.

The ion exchanger is a pretreatment system which acts as a method to fight membrane scaling in reverse osmosis. The ion exchanger works by using sodium carboxylate as a resin which will react with Chlorine and Magnesium ions, removing them from the feed. The sodium is substituted out and enters the feed stream, increasing the overall sodium level and TDS. The lack of magnesium and chlorine ions means that there is nothing to react with the carbonates to form scaling salts on the membranes.

==Desalination (Reverse Osmosis)==

Two pumps (P-105 & P-106) take the adjusted feed seawater up to 58 bar. However, before this pumping occurs, about 70% of the original stream gets separated and heads into a PX Pressure Exchanger. This separated stream later joins together with the pumped stream to enter the reverse osmosis system at the required pressure of 58 bar. The streams can combine together as they have equal pressure. It is the concentrate stream from the output of the reverse osmosis system that transfers its energy to the adjusted feed in the PX Pressure Exchanger.

[[File:Figure_3.jpg|left|300px|frame|Figure 3. Inside look of pressure exchanger]]
The pressure of the concentrate stream is about 57 bar and it is fed into an exchanger which works by using vessel that resembles a small pipe within a rotating wheel. The low pressure (LP) feed enters the vessel at 1 bar (see Figure 3). The feed turns and is pushed out into an outlet pipe by the high pressure (HP) concentrate stream at ~57 bar. The feed is now a HP stream via this exchanged energy. The now LP concentrate stream turns and is pushed out into an outlet pipe by the LP feed entering the vessel.

98% of the energy within the concentrate stream can be recovered via the PX Pressure Exchanger.<ref>Desalination Product and Applications Catalog 2013. “Advance Technologies for Seawater Reverse Osmosis and Brackish Water Applications”. Energy Recovery, Inc. 2012-2014.</ref >Thus, the now HP feed stream is not quite 58 bar and will need to be pumped by a small booster pump (P-107).

When designing a reverse osmosis system, we need to make decisions on what specific kind of reverse osmosis system will be implemented. When making the decision, we investigated the advantages of different types of reverse osmosis among the following options: batch vs. continuous process, number of stages, and number of elements.<ref>FILMTEC Reverse Osmosis Membranes. “Water & Process Solutions Technical Manual.” Dow Chemical Company.</ref> Our decision was made based on process optimization which is described later in this report. The end result of that optimization was to go with a single stage, continuous process with five elements in series.

=Post Treatment=

The permeate is sent to post treatment after the salt content is analyzed and controlled if too high or too low due to any unexpected errors. As stated above, the post treatment process will be done 50 meters from our plant by the Long Beach Water Reclamation Plant. The reason we aren’t performing post treatment “in-house” is due to the increased cost it would take for us to build and run the equipment for post treatment. We have the option to use the government’s water treatment facilities that already exist and our 15 million gallons per day output is a relatively small addition to the large Long Beach Water Reclamation Plant. We will assume any costs to post treatment will be pushed to the consumer in the price of purchasing the drinking water from the Long Beach Water Reclamation Plant.

Post treatment at any facility works to bring water to “taste standards”. The water that is the permeate of our process is already meeting L.A. drinking standards as shown above in Table 1 after our mass balance calculations around the reverse osmosis system (Appendix 3, Table 1). The post treatment increases the pH from 6.3 to 7.3 with the addition of base and then adds minerals like fluoride, calcium, and magnesium back in to improve taste.

==Brine Treatment and Byproduct Waste==

A natural consequence of a reverse osmosis procedure is that a lot of salt will be leftover from the membrane separation in a very concentrated form. This is called “brine” and will have to be taken care of. According to EPA requirements, the salt concentration of water entering the environment (via land or sea) cannot be higher than 40,000 ppm.8 The salt concentration of our brine is much higher (around 70,000 ppm). Thus, the brine must be removed or treated. We decided to treat the brine by diluting it as removing brine can be difficult and may require lots of land (e.g. evaporation ponds).

After the concentrate exchanges energy in the energy exchanger, it becomes a low pressure stream. It is pumped into a dilution tank and mixes with 18.2 million gallons per day of seawater from the inlet to become diluted brine which is pumped 600 meters out to sea.

The waste byproducts are the solid buildup on the filters and membranes. The filters and membranes also have a set lifetime and thus will also need to be removed. The waste is all non-hazardous and thus can be put in the municipal dump.

=Process Details and Equipment Sizing=

==Filters and UV Treatment [Pretreatment]==
The cost of our ultrafiltration system is estimated using data from the University of Tennessee at Knoxville.<ref>Qiang He and R. Bruce Robinson Membrane Filtration Systems. “Membrane Module Costs.” University of Tennessee at Knoxville. December 2014.</ref> It estimated the cost of the cost of the ultrafiltration $0.057 per 1000 gallons and the construction and capital costs are $0.39 per gallon per day. The cost of UV treatment is based off of units available from buyultraviolet.com for units that can treat about 450 gallons per minute.<ref>Atlantic Ultraviolet Corporation. “Megatron Water Disinfection System.”</ref> Due to the water having to be in relatively small pipes for treatment we do not expect the costs to scale very much, so we believe these units to be an accurate cost estimate.

==Ion Exchanger [Pretreatment] ==

The ion exchanger was sized using data from DOW. Our process will use 60 vessels with ion exchange beds of volume 5.73 m3. With a flow rate of 30 bed volumes/hour (10,312.2 m3/hr) the system of beds has a capacity of 125% of the feed. This allows for periodic regeneration of the resin, a weak acid cation resin (DOW MAC-3 Resin). This weak cation resin will remove cationic calcium and magnesium from the water to prevent scaling. The mass flow was calculated via a mass balance on the ion exchanger (Appendix 3).

==Reverse Osmosis System [Desalination]==

Each of the 2607 Reverse Osmosis SW30HRLE-440i [Seawater 30 High Rejection Low Energy 440 ft2 interlocking cap membrane] units have the same dimensions. A, the net length, is 1016 mm. B, the total length including the endcap, is 1029 mm. C, the diameter of the filter element, is 201mm. D, the diameter of the permeate pathway, is 29 mm. For a single stage, we have five of the elements in series so the overall length of a single stage is expected to be roughly 5145 mm. The membrane material is polyamide thin film composite. It will last three years.
Operating Conditions [ROSA 9.0]: 
* Temperature: 30oC 
* Pressure: 58 bar 
* Flow Rate: 52.3 Million gal per day 
* Recovery: 28.7%

[[File:Figure_4.png|center|400px|frame|Figure 4. Diagram of Dow FILMTEC SW30HRLE-440i]]
Cleaning Schedule: 
* Conservative Estimate: Once every three months (three days available for cleaning)<ref name = "Bates"/> 
* DOW FILMTEC Cleaning Conditions 
** The normalized permeate flow drops 10%
* The normalized salt passage increases 5 - 10%

==Seawater Holding Tank and Dilution Tank [Brine Treatment]==

The initial seawater holding tank works as a buffer to prevent surges and help perform shutdowns quickly and easily. We expect to run continuously for 350 days a year with the rest of the available time spread out and used for shutting down the plant for cleaning membranes, filters, and the fixing of unexpected errors. There is a big tradeoff between having a holding tank with high retention time and having a holding tank with low cost. With that in mind, we erred on the side of lowering capital cost and went with a retention time of one hour. The volume could be found by multiplying the retention time with the feed flow rate (Stream 1). We assumed that a good height to width ratio for a large water tank would be 4:1.

A mass balance was done around the dilution tank to find what amount of seawater would be necessary to dilute the brine to 45,000 ppm salinity (Appendix 3; Table 3). With this the feed flow rate into the tank could be found. It was found that it takes about 20 minutes of retention time for good dilution of the brine in seawater.8 We again assumed that a good height to width ratio for a large water tank would be 4:1.

We assumed a corrosion allowance of 4 mm wall thickness. We calculated wall thickness using vessel pressure, hoop stress, longitudinal stress, and weld factor. Each tank is made with SS316 to protect against corrosion.<ref>Towler G, Sinnott R. Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013.</ref>

<center>
'''Table 2.''' Tank sizes.
[[File:Table_2.jpg|center|600px]]
</center>

=Process Optimization=
<center> '''Table 3.'''Optimization Criteria
[[File:Table_3.jpg|center|400px]]
</center>
In order to optimize the reverse osmosis system, numerous simulations were run. Three representative cases are reported as shown above (Table 3). The objective of the optimization was to minimize the required water price to break even in 10 years. In order to achieve this goal, the balance between the capital cost from the number of the filter elements and utility cost from energy consumption was weighed. Before the energy recovery system was implemented, the general trend was that if we have more filter elements in series, overall recovery increases but recovery rate per filter decreases, and we need more filter elements. While increasing the elements in series increase the required number of filters, and thereby increases the capital cost, we require less energy to produce the same amount of water. After the energy recovery system is implemented, the correlation between the number of elements in series and the required energy becomes more complicated because high pressure concentrate is used to pressurize the feed. However, the trade-off between the utility cost and capital cost still exists. Of the three cases shown above, single stage, five elements RO system turned out to be the most optimal solution because single stage, single element system required higher utility cost and two stage system slightly did not meet the LA standard. Detailed information which was used for the optimization is available in Appendix 3 and in Appendix 6.

=Ethical, Environmental, and Societal Considerations=

There are a few key considerations we have followed to meet ethical, environmental, and societal guidelines and requirements: 
* Meet L.A. Drinking Water Standards
** Correct ion concentration
* Moved near water treatment plant
** Pipes will not corrode from low pH
* Working below pressure limits
* Water will be locally served and meet taste standards
** Via Post Treatment mineralization
* Follows EPA Environmental Salinity Requirements
* Reduced Energy Requirements
** PX Pressure Exchanger
** Optimized RO System
** Reverse Osmosis is itself Energy Saving

For further environmental considerations, see effluent summary in Appendix 4.

=Economic Analysis=

The pricing of the majority of the process was completed using the Aspen Economic Analyzer, however we had determine costs for the ultrafiltration, reverse osmosis membranes, and the ultraviolet treatment was completed separately. This led to a total ISBL of $56 MM, OSBL of $22.4 MM, Engineering Costs of $22.4 MM and a Contingency of $7.8 MM. We also chose a working Capital of $10 MM. This working capital is sufficient because we have no raw materials costs. We also had FCOP of $5.2 MM and VCOP of $3.5 MM.

We originally were planning on pricing our water at $700 per acre-foot, but after doing economic analysis we had a 10 year NPV of $-76.2 MM. After determining did a sensitivity analysis and determined that if our water was priced at $2104 per acre-foot we could have an NPV of $0 after 10 years and if we were able to get the same price of $2257 per acre-foot as the plant in Carlsbad we would have a 10 year NPV of $8.3 MM which increases to $41.1 MM at 20 years. For more detailed data about the economic analysis see Appendix 6.

=Conclusion and Recommendations=

Our mission was to create an economically viable process that will desalinate water for human consumption in the L.A. County. We have found after research and optimization that the best process to undergo desalination is a continuous single stage, five elements reverse osmosis system. Our process can meet the required L.A. drinking water standard and we have set up plans for post treatment at a local water reclamation site. While meeting the drinking water standards, we also met our ethical, environmental, and social standards.

Nevertheless, at a sale price of $700 per acre-foot, our proposed desalination plant in Los Angeles, California looks to be a very expensive project to undergo. Without any sort of government assistance to carry the heavy capital costs of such a project, we would suggest the plant not be built. However, after performing a sensitivity analysis at higher (and less conservative) price points we reached a breakeven point in 10 years with a sale price of $2104 per acre-foot of drinking water. Research has shown that the droughts striking California look only to get worse in the near future. With limited freshwater resources, there is no option other than to consider desalination. Therefore it is reasonable to expect we could get near the same price deal as local desalination plants like the Carlsbad Plant ($2257 per acre-foot) As we have met all design constraints for this project, we hope to convince policymakers to adopt our project at least $2104 per acre-foot water. If this deal is reached, we recommend that this project and design be funded.

However, there are ways to improve our design with future work. Right now all the pipes diameters are based on having a flowrate of 1.5 m/s. But in reality, we’d need to use standard piping sizes to get more accurate pricing and energy costs. Also, in reality we would be purchasing some of these vessels in bulk quantities (i.e. bulk filters) so the price could be lowered with negotiation and economies of scale. We also have pumps that are providing very little pressure for their high cost. We can remove these pumps and simply run at higher pressures earlier. We can potentially reduce early energy costs by building a “floating” desalination plant in the sea to reduce pumping distance from the well.

=Appendices=

==Appendix 1: Design Basis==
[[File:Appendix_1.png|center|Appendix_1.png]]

==Appendix 2: Process Flow Diagram==
[[File:Appendix_2.png|center]]
==Appendix 3: Mass and Energy Balance Tables==
'''Table 1'''. Mass balance on 1 stage 5 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T1.png|center|500px|]]

'''Table 2'''. Mass balance on 1 stage 1 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T2.png|center|500px|]]

'''Table 3'''. Mass balance on 2 stage 5,5 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T3.png|center|600px|]]

'''Table 4'''. 1 stage 5 elements CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T4.png|center|300px|]]

'''Table 5'''. 1 stage 1 element CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T5.png|center|300px|]]

'''Table 6'''. 2 stages 5,5 elements CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T6.png|center|300px|]]

'''Table 7'''. Energy balance on 1 stage 5 elements RO system
[[File:Appendix_3_T7.png|center|800px|]]

==Appendix 4: Effluent Summary==
[[File:Appendix_4.png|center]]

==Appendix 5: Cost of Production==
[[File:Appendix_5.png|center]]

==Appendix 6: Economic Anlaysis==
===Case I ===
Water is $700; RO with 1 stage and 5 elements
[[File:Appendix_6_C1.png|center]]

===Case II===
Water is $1400; RO with 1 stage and 5 elements
[[File:Appendix_6_C2.png|center]]

===Case III: Carlsbad ===
Water is $2257; RO with 1 stage and 5 elements
[[File:Appendix_6_C3.png|center]]

===Case IV: Breakeven (NPV = 0 at 10 years)===
Water is $2104; RO with 1 stage and 5 elements
[[File:Appendix_6_C4.png|center]]

===Case V: Breakeven (NPV = 0 at 10 years)===
Water is $2112; RO with 2 stages and 5,5 elements
[[File:Appendix_6_C5.png|center]]

===Case VI: Breakeven (NPV = 0 at 10 years)===
Water is $2138 with 1 stage and 1 element
[[File:Appendix_6_C6.png|center]]

==Appendix 7: ROSA Simulation Outputs==

'''Table 1'''. System detail of 1 stage 5 elements RO system
[[File:Appendix_7_T1.png|center|500px|]]

'''Table 2'''. System detail of 1 stage 1 elements RO system
[[File:Appendix_7_T2.png|center]]

'''Table 3'''. System detail of 2 stage 5,5 elements system
[[File:Appendix_7_T3.png|center]]

'''Table 4'''. 1 stage 5 elements RO system concentrations and pH in each stream
[[File:Appendix_7_T4.png|center]]

'''Table 5'''. 1 stage 1 element RO system concentration and pH in each stream
[[File:Appendix_7_T5.png|center]]

'''Table 6'''. 2 stage 5,5 elements RO system concentrations and pH in each stream
[[File:Appendix_7_T6.png|center]]

=References=
<references/>

Desalination - Team E

2016-03-12T03:37:19Z

Bts511: /* Appendix 7: ROSA Simulation Outputs */

Group E Corporation

Authors: Hassan Ali, Woo Soo Choe, Brett Sleyster, Jake Stolley

Instructors: Fengqi You, David Wegerer

March 11, 2016

=Executive Summary=

Outlined in this report is the project model report for the proposed desalination plant in Los Angeles, California which will provide drinking water to local citizens using reverse osmosis process. This location was chosen due to the region’s propensity for droughts in recent years. The fresh water resources in California have plummeted as water prices have risen 3-4% each year. This is only expected to continue. Other companies have worked deals with the Californian government to create desalination plants that provide a new fresh water drinking source. This suggests that there is a strong market here with potential growth.

The plant will have an output of 15 million gallons per day, which is about medium size relative to other plants. Thus the build time will be about one year. The plant will run for 30 years, with 350 days per year of use. Shutdown time will be used for cleaning and maintenance of equipment. The plant will be built on Long Beach, 500 meters from the coast nearby the Long Beach Water Reclamation Plant (LBWP). The input will withdraw 70.5 million gallons per day of seawater that will be used for reverse osmosis and dilution of brine.

Reverse osmosis was chosen as the desalination method for its low energy cost and flexibility. We designed a process that will pretreat the seawater to remove waste, colloids, bacteria, and membrane-damaging scalants. Once this seawater is treated is pumped to heavy pressures through a reverse osmosis system. After doing process optimization under various system parameter changes, we found a continuous, single stage, five element reverse osmosis process to be best. It produces a near 30% recovery. We reduce energy costs by including a pressure exchanger between our concentrate outlet and adjusted feed into the reverse osmosis system.

After performing mass and energy balances, we were able to size equipment and find our permeate composition. It turns out our process will meet L.A. standards for drinking water quality. However, post treatment will be done by the local LBWP to make the water taste better and be non-caustic to pipes. This will reduce “in-house” costs. Many environmental considerations were made, to meet our environmental standards. We meet EPA salinity requirements for our brine waste by diluting it. We also don’t use a lot of energy and thus reduce our fossil fuel consumption and CO2 generation. To be safe, we work well under max pressure limits for our equipment. We make sure we are ethical and meet all standards that our local county requires of us.

After an economic evaluation at a conservative well-water price of $700 per acre-foot, we realize that our revenue is way too low to cover our costs. Capital costs are especially large with regards to our large water tanks. We find that we reach a 10 year NPV breakpoint of $2104 per acre-foot after performing a sensitivity analysis. The Carlsbad Plant in San Diego worked out a deal for $2257 per acre-foot.

As we have met all design constraints for this project, we hope to convince policymakers to adopt our project at least $2104 per acre-foot water. If this deal is reached, we recommend that this project and design be funded. There is future work to be done to improve costs such as performing pump optimization. However, this work is minimal and won’t harm our bottom line if we reach our proposed price points.

__TOC__

=Introduction=

Due to a combination of various factors--including rising human population, climate change, increased energy consumption, and land erosion--the availability of fresh water as a resource for human consumption has been dwindling. This has had a negative impact on the world, but this impact has been especially felt in areas with natural shortages in fresh water. Desert regions, such as Saudi Arabia, have resorted to desalinating seawater at the coast and transporting up to their cities. This encompasses their entire freshwater system as very few natural water sources remain. In the United States there is a region which has shown increasing signs of following in the footsteps of Saudi Arabia and becoming fully reliant on desalination: Los Angeles, California.

=Location and Market Analysis=

In this project, Los Angeles was chosen as the location our plant, because the area has been subject to a severe drought for the last four years, as shown in Figure 1. The rapid decrease in the available water has caused permanent damages to the ground, decreased the amount of water in the reservoir from 45% to 25%, and reduced the amount of hydroelectric power produced in the state. Such deficiency in usable water has increased the water prices in Los Angeles by 6-7% annually for the last decade and is expected to rise by 3.4% annually for typical customers for the next five years. For customers who use more than 20000 gallons a year, the bill is expected to increase by 34% by 2021. This price increase is occurring because of both drought and maintenance. <ref>Pacific Institute: The California Drought. 2014.</ref>

[[File:Figure_1.jpg|right|200px|frame|Figure 1. California Drought Index.]]
As of March 2015 the Metropolitan Water District of Southern California was willing to pay as much as $700 per acre-foot of water to farmers in the Sacramento Valley, so we assume this will be the target price of the water produced by the this desalination plant.<ref>Hessel, Phil. “Dry Southern California Offers Northern Farmers Top Dollar for Water”. NBC News. March 18, 2015. </ref>

Currently, there is a plant in Carlsbad which supplies drinking water to the San Diego area. The plant can produce up to 50 million gallons per day and is the nation's largest water desalination plant. The water from this plant costs $2257 per acre-foot for the first 48000 acre-feet and about $2000 for any additional acre-foot. The county signed a contract with this plant to purchase the first 48000 acre-feet of water for the next 30 years.<ref>Carlsbad Desalination Plant. “Project Overview”. January 2016.</ref> Also, there is another desalination plant in Tampa that produces about 25 million gallons of water per day, and over 2000 other public and industrial plants with production capacity of over 300000 gallons per day. When conducting the economic analysis our plant, such numbers presented by other plants must be considered, so our project can yield positive profit.<ref>Leven, Rachel. “U.S. Desalination Industry Grows Since 2000; See as Essential to Meeting Supply Needs”. Bloomberg BNA. August 21, 2013.</ref>

The profitability of the project is naturally heavily dependent on the price of water in California. If the drought ends, water supply will increase and the water price will drop significantly, but if the drought continues, the price of water in California could rise even more rapidly and allow desalination plants to be more profitable.<ref>Southern California Public Radio. “Where is California water use decreasing?”. December 2015.</ref> It may be difficult to make this plant profitable without a guaranteed purchase or other subsidy from the local governments because the presence of pre-existing plants maintains the water price relatively low. However, external sources of profit such as guaranteed purchase or subsidy are expected because the ramifications of not having desalination plant or other sources of water may be significantly more expensive than the subsidizing a desalination plant assuming the drought continues.<ref>Weiser, Matt. “Could desalination solve California’s water problem?”. The Sacramento Bee. October 18, 2014.</ref>

=Design Basis and Technical Approach=

[[File:Figure_2.jpg|left|200px|frame|Figure 2. Vertical Well Feed]]
Considering the number of pre-existing desalination plant and the potential risk of the drought ending, our plant will be tentatively designed to produce 15 million gallons of drinkable water per day to homes and businesses in the Los Angeles county area. With a recovery rate of near 30%, and a requirement for diluting any brine, the input feed flowrate will be 70.5 million gallons per day. Following the similarly built Carlsbad plant, we expect the plant to last 30 years. The project building timeline is one year based on our relatively small plant size. We will run for 350 days a year while shutting down for cleaning and maintenance once every 3 months (based on DOW standards).<ref name = "Bates">Bates, Wayne. “Cleaning your RO”. Hydranautics. </ref> The plant will be built 500 meters from the coastline on Long Beach. A borehole will be drilled 10 meters into the ground. At this depth there should be seawater aquifers below the coast that regularly recharges itself with seawater from the ocean. A vertical well containing intake holes will be put into the borehole. The intake holes will be equipped with large size filters that can easily be replaced. At the surface, the well will be connected to a large-scale pump in a well pump house. From there, the seawater will be horizontally pumped into a seawater tank in the desalination plant. This will be considered the start of our process. The process ends after pretreatment, desalination, post treatment, and brine treatment are completed. Appendix 1 shows more details on the Design Basis.

The feed composition is assumed to be the standard seawater “average” based on the DOW Chemical ROSA 9.0 Technical Manual in Table 1. However, in reality the seawater composition at the plant site will have to be chemically tested. The final product will be based on L.A. standards for drinking water quality (Table 1). The specifications in the table are upper limits that we cannot exceed. We will aim to reach the salt level standards but any remineralization and pH treatment will be exported and paid for due to the added complexity and cost in making such a system in-house.

Due to the acidity of the desalinated water coming out of the reverse osmosis (RO) system (pH ~ 6), the pumps and pipes carrying the desalinated water for post treatment will likely corrode over time. We have decided to place the desalination plant by the Long Beach Water Reclamation Plant (about 50 meters away) to prevent pipe corrosion.

[[File:Table_1.jpg|center|200px|frame|Table 1. Concentration of the standard seawater, the output from the RO system, and LA standard.]]

=Desalination Overview and Alternatives=

We first began our research into the current landscape of the desalination industry, particularly of plants that were built in the United States. We discovered there were many processes but the main ones included: 
* Membrane Separation (Forward Osmosis and Reverse Osmosis)
* Thermal Separation (Multi Stage Flash Distillation)
* Phase Change Separation (Freeze Separation)
Each process was analyzed for their pros and cons. For our project design, we decided to go with reverse osmosis as our desalination process. This is because it is the most commonly used process in the industry, and for good reason. Unlike thermal separation, which requires huge amounts of energy to supply heat, reverse osmosis is very energy efficient. This reduces energy costs that would be spread across the plant’s lifetime.<ref name = "Republic">Republic of Mauritius NY. "Technical Aspects of Desalination Plant". Sustainable Sanitation and Water Management. </ref>

Forward osmosis is a newer technology being developed that also uses membranes to separate water and salt and may be more energy efficient overall than reverse osmosis. This is because forward osmosis doesn’t require a pump to pressurize seawater through the membranes, but rather leaves a higher salt concentrated solution (AKA draw solution) on the other side of the membrane for the seawater to be drawn into. Then a second step is required to regenerate the concentrated draw solution and produce purified water.<ref>Mallinson, Alissa. "Study Shows Forward Osmosis Desalination Not Energy Efficient." MIT News [Boston]. July 23, 2014.</ref> We decided to move away from forward osmosis because it is a process that is still being refined and developed for desalination and lacks the data, programs, and technologies that reverse osmosis users benefit from. Additionally, a recent study from MIT has claimed that due to the inherent nature of forward osmosis requiring that seawater be drawn to a higher concentrated salt solution before the water and salt is separated, it cannot possibly be more energy efficient than reverse osmosis.<ref>O'Hern, Sean C., Karnik, Rohit. Nanofiltration across Defect-Sealed Nanoporous Monolayer Graphene. American Chemical Society 15.5 (2014); 3254-260. </ref>

Other processes such as freeze separation were too restrictive in their feed and process demands and required a lot of capital. Ultimately, reverse osmosis is simply the best process that exists for desalination and has been ranked the best when performance is evaluated in terms of technical, environmental, and economic aspects. It is a process that is flexible in water quantity, operation start-up and shut-off, and site location.

Nevertheless there are disadvantages to reverse osmosis that must be considered. The feed seawater must not be too salty and too low quality. The pressure to pump the seawater is enormous. The construction time is long for large plants. The capital costs are also high when compared with other processes such as thermal separation. But all of these disadvantages are ultimately outweighed by the advantages reverse osmosis provides.<ref name = "Republic"/>

=Process Overview and Alternatives=

See Appendix 2 (Figure 1) for the Process Flow Diagram (PFD). The process itself is detailed below. The stream flow rates are in Appendix 2 (Table 1). We assumed 10 meters of pipe for every stream except the inlet and outlet streams of the entire PFD.

==Pretreatment==

After the 10 meter vertical well pumps seawater from the Long Beach coast, the feed is pumped 500 meters to a seawater tank that makes it easy to shut-down the system and reduce the large volume of flow. The level control keeps the seawater from overfilling the tank. From there, part of the seawater goes to brine treatment while the rest go to pretreatment. The purpose of pretreatment in the desalination process is to prevent equipment damage by removing particles and microorganisms from the water. A variety of pollutants -- including calcium carbonate, sulfates, silica, clay, bacteria, and other colloids and microorganisms -- can disrupt the reverse osmosis process. Media filtration uses fine grains such as sand or anthracite to filter the particles. We used a large filter to catch larger particle material (e.g. sand). However, simple filtration is not enough to prevent colloid fouling. An ultrafilter (i.e. microfiltration) is used to remove virtually all suspended matter, effectively preventing colloid fouling. The costs associated with ultrafiltration are not prohibitive in terms of the project’s budget.<ref name = "Desalination">Desalination Pre-treatment. Lenntech.</ref>

Reverse osmosis systems can also be disrupted by the growth of microbiological films of bacteria on the film and pipe walls. They are typically treated as colloids during pretreatment, and most organisms are removed from the water by ultrafiltration, but this does not remove the smallest of organic particles. DBNPA is the most effective biocide, but these aren’t often used and have to be removed from the water. Chlorination is an effective way to remove bacteria, but it can cause degradation of the membrane, and must be removed before the reverse osmosis occurs. In our process we will use UV irradiation to remove any organic material that passes through the microfilters.<ref name = "Desalination"/>

Scaling can reduce the flux across the membrane and can build up in pipe corners, causing backup. It is caused by inorganic salts -- specifically sulfates and calcium and magnesium carbonate -- present in the water. Scaling can be prevented and/or controlled with acids, antiscalants, or ion exchange softening. Antiscalants are good for removing sulfates, but not much else. Acidification will only be effective at removing CaCO3 from water but, the decreased pH may deteriorate the membrane and pipes, so it may require future research.<ref name = "Desalination"/> Ion exchange is very expensive and requires more capital to undergo, but is overall the more effective option.

The ion exchanger is a pretreatment system which acts as a method to fight membrane scaling in reverse osmosis. The ion exchanger works by using sodium carboxylate as a resin which will react with Chlorine and Magnesium ions, removing them from the feed. The sodium is substituted out and enters the feed stream, increasing the overall sodium level and TDS. The lack of magnesium and chlorine ions means that there is nothing to react with the carbonates to form scaling salts on the membranes.

==Desalination (Reverse Osmosis)==

Two pumps (P-105 & P-106) take the adjusted feed seawater up to 58 bar. However, before this pumping occurs, about 70% of the original stream gets separated and heads into a PX Pressure Exchanger. This separated stream later joins together with the pumped stream to enter the reverse osmosis system at the required pressure of 58 bar. The streams can combine together as they have equal pressure. It is the concentrate stream from the output of the reverse osmosis system that transfers its energy to the adjusted feed in the PX Pressure Exchanger.

[[File:Figure_3.jpg|left|300px|frame|Figure 3. Inside look of pressure exchanger]]
The pressure of the concentrate stream is about 57 bar and it is fed into an exchanger which works by using vessel that resembles a small pipe within a rotating wheel. The low pressure (LP) feed enters the vessel at 1 bar (see Figure 3). The feed turns and is pushed out into an outlet pipe by the high pressure (HP) concentrate stream at ~57 bar. The feed is now a HP stream via this exchanged energy. The now LP concentrate stream turns and is pushed out into an outlet pipe by the LP feed entering the vessel.

98% of the energy within the concentrate stream can be recovered via the PX Pressure Exchanger.<ref>Desalination Product and Applications Catalog 2013. “Advance Technologies for Seawater Reverse Osmosis and Brackish Water Applications”. Energy Recovery, Inc. 2012-2014.</ref >Thus, the now HP feed stream is not quite 58 bar and will need to be pumped by a small booster pump (P-107).

When designing a reverse osmosis system, we need to make decisions on what specific kind of reverse osmosis system will be implemented. When making the decision, we investigated the advantages of different types of reverse osmosis among the following options: batch vs. continuous process, number of stages, and number of elements.<ref>FILMTEC Reverse Osmosis Membranes. “Water & Process Solutions Technical Manual.” Dow Chemical Company.</ref> Our decision was made based on process optimization which is described later in this report. The end result of that optimization was to go with a single stage, continuous process with five elements in series.

=Post Treatment=

The permeate is sent to post treatment after the salt content is analyzed and controlled if too high or too low due to any unexpected errors. As stated above, the post treatment process will be done 50 meters from our plant by the Long Beach Water Reclamation Plant. The reason we aren’t performing post treatment “in-house” is due to the increased cost it would take for us to build and run the equipment for post treatment. We have the option to use the government’s water treatment facilities that already exist and our 15 million gallons per day output is a relatively small addition to the large Long Beach Water Reclamation Plant. We will assume any costs to post treatment will be pushed to the consumer in the price of purchasing the drinking water from the Long Beach Water Reclamation Plant.

Post treatment at any facility works to bring water to “taste standards”. The water that is the permeate of our process is already meeting L.A. drinking standards as shown above in Table 1 after our mass balance calculations around the reverse osmosis system (Appendix 3, Table 1). The post treatment increases the pH from 6.3 to 7.3 with the addition of base and then adds minerals like fluoride, calcium, and magnesium back in to improve taste.

==Brine Treatment and Byproduct Waste==

A natural consequence of a reverse osmosis procedure is that a lot of salt will be leftover from the membrane separation in a very concentrated form. This is called “brine” and will have to be taken care of. According to EPA requirements, the salt concentration of water entering the environment (via land or sea) cannot be higher than 40,000 ppm.8 The salt concentration of our brine is much higher (around 70,000 ppm). Thus, the brine must be removed or treated. We decided to treat the brine by diluting it as removing brine can be difficult and may require lots of land (e.g. evaporation ponds).

After the concentrate exchanges energy in the energy exchanger, it becomes a low pressure stream. It is pumped into a dilution tank and mixes with 18.2 million gallons per day of seawater from the inlet to become diluted brine which is pumped 600 meters out to sea.

The waste byproducts are the solid buildup on the filters and membranes. The filters and membranes also have a set lifetime and thus will also need to be removed. The waste is all non-hazardous and thus can be put in the municipal dump.

=Process Details and Equipment Sizing=

==Filters and UV Treatment [Pretreatment]==
The cost of our ultrafiltration system is estimated using data from the University of Tennessee at Knoxville.<ref>Qiang He and R. Bruce Robinson Membrane Filtration Systems. “Membrane Module Costs.” University of Tennessee at Knoxville. December 2014.</ref> It estimated the cost of the cost of the ultrafiltration $0.057 per 1000 gallons and the construction and capital costs are $0.39 per gallon per day. The cost of UV treatment is based off of units available from buyultraviolet.com for units that can treat about 450 gallons per minute.<ref>Atlantic Ultraviolet Corporation. “Megatron Water Disinfection System.”</ref> Due to the water having to be in relatively small pipes for treatment we do not expect the costs to scale very much, so we believe these units to be an accurate cost estimate.

==Ion Exchanger [Pretreatment] ==

The ion exchanger was sized using data from DOW. Our process will use 60 vessels with ion exchange beds of volume 5.73 m3. With a flow rate of 30 bed volumes/hour (10,312.2 m3/hr) the system of beds has a capacity of 125% of the feed. This allows for periodic regeneration of the resin, a weak acid cation resin (DOW MAC-3 Resin). This weak cation resin will remove cationic calcium and magnesium from the water to prevent scaling. The mass flow was calculated via a mass balance on the ion exchanger (Appendix 3).

==Reverse Osmosis System [Desalination]==

Each of the 2607 Reverse Osmosis SW30HRLE-440i [Seawater 30 High Rejection Low Energy 440 ft2 interlocking cap membrane] units have the same dimensions. A, the net length, is 1016 mm. B, the total length including the endcap, is 1029 mm. C, the diameter of the filter element, is 201mm. D, the diameter of the permeate pathway, is 29 mm. For a single stage, we have five of the elements in series so the overall length of a single stage is expected to be roughly 5145 mm. The membrane material is polyamide thin film composite. It will last three years.
Operating Conditions [ROSA 9.0]: 
* Temperature: 30oC 
* Pressure: 58 bar 
* Flow Rate: 52.3 Million gal per day 
* Recovery: 28.7%

[[File:Figure_4.png|center|400px|frame|Figure 4. Diagram of Dow FILMTEC SW30HRLE-440i]]
Cleaning Schedule: 
* Conservative Estimate: Once every three months (three days available for cleaning)<ref name = "Bates"/> 
* DOW FILMTEC Cleaning Conditions 
** The normalized permeate flow drops 10%
* The normalized salt passage increases 5 - 10%

==Seawater Holding Tank and Dilution Tank [Brine Treatment]==

The initial seawater holding tank works as a buffer to prevent surges and help perform shutdowns quickly and easily. We expect to run continuously for 350 days a year with the rest of the available time spread out and used for shutting down the plant for cleaning membranes, filters, and the fixing of unexpected errors. There is a big tradeoff between having a holding tank with high retention time and having a holding tank with low cost. With that in mind, we erred on the side of lowering capital cost and went with a retention time of one hour. The volume could be found by multiplying the retention time with the feed flow rate (Stream 1). We assumed that a good height to width ratio for a large water tank would be 4:1.

A mass balance was done around the dilution tank to find what amount of seawater would be necessary to dilute the brine to 45,000 ppm salinity (Appendix 3; Table 3). With this the feed flow rate into the tank could be found. It was found that it takes about 20 minutes of retention time for good dilution of the brine in seawater.8 We again assumed that a good height to width ratio for a large water tank would be 4:1.

We assumed a corrosion allowance of 4 mm wall thickness. We calculated wall thickness using vessel pressure, hoop stress, longitudinal stress, and weld factor. Each tank is made with SS316 to protect against corrosion.<ref>Towler G, Sinnott R. Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013.</ref>

<center>
'''Table 2.''' Tank sizes.
[[File:Table_2.jpg|center|600px]]
</center>

=Process Optimization=
<center> '''Table 3.'''Optimization Criteria
[[File:Table_3.jpg|center|400px]]
</center>
In order to optimize the reverse osmosis system, numerous simulations were run. Three representative cases are reported as shown above (Table 3). The objective of the optimization was to minimize the required water price to break even in 10 years. In order to achieve this goal, the balance between the capital cost from the number of the filter elements and utility cost from energy consumption was weighed. Before the energy recovery system was implemented, the general trend was that if we have more filter elements in series, overall recovery increases but recovery rate per filter decreases, and we need more filter elements. While increasing the elements in series increase the required number of filters, and thereby increases the capital cost, we require less energy to produce the same amount of water. After the energy recovery system is implemented, the correlation between the number of elements in series and the required energy becomes more complicated because high pressure concentrate is used to pressurize the feed. However, the trade-off between the utility cost and capital cost still exists. Of the three cases shown above, single stage, five elements RO system turned out to be the most optimal solution because single stage, single element system required higher utility cost and two stage system slightly did not meet the LA standard. Detailed information which was used for the optimization is available in Appendix 3 and in Appendix 6.

=Ethical, Environmental, and Societal Considerations=

There are a few key considerations we have followed to meet ethical, environmental, and societal guidelines and requirements: 
* Meet L.A. Drinking Water Standards
** Correct ion concentration
* Moved near water treatment plant
** Pipes will not corrode from low pH
* Working below pressure limits
* Water will be locally served and meet taste standards
** Via Post Treatment mineralization
* Follows EPA Environmental Salinity Requirements
* Reduced Energy Requirements
** PX Pressure Exchanger
** Optimized RO System
** Reverse Osmosis is itself Energy Saving

For further environmental considerations, see effluent summary in Appendix 4.

=Economic Analysis=

The pricing of the majority of the process was completed using the Aspen Economic Analyzer, however we had determine costs for the ultrafiltration, reverse osmosis membranes, and the ultraviolet treatment was completed separately. This led to a total ISBL of $56 MM, OSBL of $22.4 MM, Engineering Costs of $22.4 MM and a Contingency of $7.8 MM. We also chose a working Capital of $10 MM. This working capital is sufficient because we have no raw materials costs. We also had FCOP of $5.2 MM and VCOP of $3.5 MM.

We originally were planning on pricing our water at $700 per acre-foot, but after doing economic analysis we had a 10 year NPV of $-76.2 MM. After determining did a sensitivity analysis and determined that if our water was priced at $2104 per acre-foot we could have an NPV of $0 after 10 years and if we were able to get the same price of $2257 per acre-foot as the plant in Carlsbad we would have a 10 year NPV of $8.3 MM which increases to $41.1 MM at 20 years. For more detailed data about the economic analysis see Appendix 6.

=Conclusion and Recommendations=

Our mission was to create an economically viable process that will desalinate water for human consumption in the L.A. County. We have found after research and optimization that the best process to undergo desalination is a continuous single stage, five elements reverse osmosis system. Our process can meet the required L.A. drinking water standard and we have set up plans for post treatment at a local water reclamation site. While meeting the drinking water standards, we also met our ethical, environmental, and social standards.

Nevertheless, at a sale price of $700 per acre-foot, our proposed desalination plant in Los Angeles, California looks to be a very expensive project to undergo. Without any sort of government assistance to carry the heavy capital costs of such a project, we would suggest the plant not be built. However, after performing a sensitivity analysis at higher (and less conservative) price points we reached a breakeven point in 10 years with a sale price of $2104 per acre-foot of drinking water. Research has shown that the droughts striking California look only to get worse in the near future. With limited freshwater resources, there is no option other than to consider desalination. Therefore it is reasonable to expect we could get near the same price deal as local desalination plants like the Carlsbad Plant ($2257 per acre-foot) As we have met all design constraints for this project, we hope to convince policymakers to adopt our project at least $2104 per acre-foot water. If this deal is reached, we recommend that this project and design be funded.

However, there are ways to improve our design with future work. Right now all the pipes diameters are based on having a flowrate of 1.5 m/s. But in reality, we’d need to use standard piping sizes to get more accurate pricing and energy costs. Also, in reality we would be purchasing some of these vessels in bulk quantities (i.e. bulk filters) so the price could be lowered with negotiation and economies of scale. We also have pumps that are providing very little pressure for their high cost. We can remove these pumps and simply run at higher pressures earlier. We can potentially reduce early energy costs by building a “floating” desalination plant in the sea to reduce pumping distance from the well.

=Appendices=

==Appendix 1: Design Basis==
[[File:Appendix_1.png|center|Appendix_1.png]]

==Appendix 2: Process Flow Diagram==
[[File:Appendix_2.png|center]]
==Appendix 3: Mass and Energy Balance Tables==
'''Table 1'''. Mass balance on 1 stage 5 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T1.png|center|500px|]]

'''Table 2'''. Mass balance on 1 stage 1 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T2.png|center|500px|]]

'''Table 3'''. Mass balance on 2 stage 5,5 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T3.png|center|600px|]]

'''Table 4'''. 1 stage 5 elements CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T4.png|center|300px|]]

'''Table 5'''. 1 stage 1 element CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T5.png|center|300px|]]

'''Table 6'''. 2 stages 5,5 elements CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T6.png|center|300px|]]

'''Table 7'''. Energy balance on 1 stage 5 elements RO system
[[File:Appendix_3_T7.png|center|800px|]]

==Appendix 4: Effluent Summary==
[[File:Appendix_4.png|center]]

==Appendix 5: Cost of Production==
[[File:Appendix_5.png|center]]

==Appendix 6: Economic Anlaysis==
===Case I ===
Water is $700; RO with 1 stage and 5 elements
[[File:Appendix_6_C1.png|center]]

===Case II===
Water is $1400; RO with 1 stage and 5 elements
[[File:Appendix_6_C2.png|center]]

===Case III: Carlsbad ===
Water is $2257; RO with 1 stage and 5 elements
[[File:Appendix_6_C3.png|center]]

===Case IV: Breakeven (NPV = 0 at 10 years)===
Water is $2104; RO with 1 stage and 5 elements
[[File:Appendix_6_C4.png|center]]

===Case V: Breakeven (NPV = 0 at 10 years)===
Water is $2112; RO with 2 stages and 5,5 elements
[[File:Appendix_6_C5.png|center]]

===Case VI: Breakeven (NPV = 0 at 10 years)===
Water is $2138 with 1 stage and 1 element
[[File:Appendix_6_C6.png|center]]

==Appendix 7: ROSA Simulation Outputs==

'''Table 1'''. System detail of 1 stage 5 elements RO system
[[File:Appendix_7_T1.png|center||600px]]

'''Table 2'''. System detail of 1 stage 1 elements RO system
[[File:Appendix_7_T2.png|center]]

'''Table 3'''. System detail of 2 stage 5,5 elements system
[[File:Appendix_7_T3.png|center]]

'''Table 4'''. 1 stage 5 elements RO system concentrations and pH in each stream
[[File:Appendix_7_T4.png|center]]

'''Table 5'''. 1 stage 1 element RO system concentration and pH in each stream
[[File:Appendix_7_T5.png|center]]

'''Table 6'''. 2 stage 5,5 elements RO system concentrations and pH in each stream
[[File:Appendix_7_T6.png|center]]

=References=
<references/>

Desalination - Team E

2016-03-12T03:36:21Z

Bts511: /* Appendix 1: Design Basis */

Group E Corporation

Authors: Hassan Ali, Woo Soo Choe, Brett Sleyster, Jake Stolley

Instructors: Fengqi You, David Wegerer

March 11, 2016

=Executive Summary=

Outlined in this report is the project model report for the proposed desalination plant in Los Angeles, California which will provide drinking water to local citizens using reverse osmosis process. This location was chosen due to the region’s propensity for droughts in recent years. The fresh water resources in California have plummeted as water prices have risen 3-4% each year. This is only expected to continue. Other companies have worked deals with the Californian government to create desalination plants that provide a new fresh water drinking source. This suggests that there is a strong market here with potential growth.

The plant will have an output of 15 million gallons per day, which is about medium size relative to other plants. Thus the build time will be about one year. The plant will run for 30 years, with 350 days per year of use. Shutdown time will be used for cleaning and maintenance of equipment. The plant will be built on Long Beach, 500 meters from the coast nearby the Long Beach Water Reclamation Plant (LBWP). The input will withdraw 70.5 million gallons per day of seawater that will be used for reverse osmosis and dilution of brine.

Reverse osmosis was chosen as the desalination method for its low energy cost and flexibility. We designed a process that will pretreat the seawater to remove waste, colloids, bacteria, and membrane-damaging scalants. Once this seawater is treated is pumped to heavy pressures through a reverse osmosis system. After doing process optimization under various system parameter changes, we found a continuous, single stage, five element reverse osmosis process to be best. It produces a near 30% recovery. We reduce energy costs by including a pressure exchanger between our concentrate outlet and adjusted feed into the reverse osmosis system.

After performing mass and energy balances, we were able to size equipment and find our permeate composition. It turns out our process will meet L.A. standards for drinking water quality. However, post treatment will be done by the local LBWP to make the water taste better and be non-caustic to pipes. This will reduce “in-house” costs. Many environmental considerations were made, to meet our environmental standards. We meet EPA salinity requirements for our brine waste by diluting it. We also don’t use a lot of energy and thus reduce our fossil fuel consumption and CO2 generation. To be safe, we work well under max pressure limits for our equipment. We make sure we are ethical and meet all standards that our local county requires of us.

After an economic evaluation at a conservative well-water price of $700 per acre-foot, we realize that our revenue is way too low to cover our costs. Capital costs are especially large with regards to our large water tanks. We find that we reach a 10 year NPV breakpoint of $2104 per acre-foot after performing a sensitivity analysis. The Carlsbad Plant in San Diego worked out a deal for $2257 per acre-foot.

As we have met all design constraints for this project, we hope to convince policymakers to adopt our project at least $2104 per acre-foot water. If this deal is reached, we recommend that this project and design be funded. There is future work to be done to improve costs such as performing pump optimization. However, this work is minimal and won’t harm our bottom line if we reach our proposed price points.

__TOC__

=Introduction=

Due to a combination of various factors--including rising human population, climate change, increased energy consumption, and land erosion--the availability of fresh water as a resource for human consumption has been dwindling. This has had a negative impact on the world, but this impact has been especially felt in areas with natural shortages in fresh water. Desert regions, such as Saudi Arabia, have resorted to desalinating seawater at the coast and transporting up to their cities. This encompasses their entire freshwater system as very few natural water sources remain. In the United States there is a region which has shown increasing signs of following in the footsteps of Saudi Arabia and becoming fully reliant on desalination: Los Angeles, California.

=Location and Market Analysis=

In this project, Los Angeles was chosen as the location our plant, because the area has been subject to a severe drought for the last four years, as shown in Figure 1. The rapid decrease in the available water has caused permanent damages to the ground, decreased the amount of water in the reservoir from 45% to 25%, and reduced the amount of hydroelectric power produced in the state. Such deficiency in usable water has increased the water prices in Los Angeles by 6-7% annually for the last decade and is expected to rise by 3.4% annually for typical customers for the next five years. For customers who use more than 20000 gallons a year, the bill is expected to increase by 34% by 2021. This price increase is occurring because of both drought and maintenance. <ref>Pacific Institute: The California Drought. 2014.</ref>

[[File:Figure_1.jpg|right|200px|frame|Figure 1. California Drought Index.]]
As of March 2015 the Metropolitan Water District of Southern California was willing to pay as much as $700 per acre-foot of water to farmers in the Sacramento Valley, so we assume this will be the target price of the water produced by the this desalination plant.<ref>Hessel, Phil. “Dry Southern California Offers Northern Farmers Top Dollar for Water”. NBC News. March 18, 2015. </ref>

Currently, there is a plant in Carlsbad which supplies drinking water to the San Diego area. The plant can produce up to 50 million gallons per day and is the nation's largest water desalination plant. The water from this plant costs $2257 per acre-foot for the first 48000 acre-feet and about $2000 for any additional acre-foot. The county signed a contract with this plant to purchase the first 48000 acre-feet of water for the next 30 years.<ref>Carlsbad Desalination Plant. “Project Overview”. January 2016.</ref> Also, there is another desalination plant in Tampa that produces about 25 million gallons of water per day, and over 2000 other public and industrial plants with production capacity of over 300000 gallons per day. When conducting the economic analysis our plant, such numbers presented by other plants must be considered, so our project can yield positive profit.<ref>Leven, Rachel. “U.S. Desalination Industry Grows Since 2000; See as Essential to Meeting Supply Needs”. Bloomberg BNA. August 21, 2013.</ref>

The profitability of the project is naturally heavily dependent on the price of water in California. If the drought ends, water supply will increase and the water price will drop significantly, but if the drought continues, the price of water in California could rise even more rapidly and allow desalination plants to be more profitable.<ref>Southern California Public Radio. “Where is California water use decreasing?”. December 2015.</ref> It may be difficult to make this plant profitable without a guaranteed purchase or other subsidy from the local governments because the presence of pre-existing plants maintains the water price relatively low. However, external sources of profit such as guaranteed purchase or subsidy are expected because the ramifications of not having desalination plant or other sources of water may be significantly more expensive than the subsidizing a desalination plant assuming the drought continues.<ref>Weiser, Matt. “Could desalination solve California’s water problem?”. The Sacramento Bee. October 18, 2014.</ref>

=Design Basis and Technical Approach=

[[File:Figure_2.jpg|left|200px|frame|Figure 2. Vertical Well Feed]]
Considering the number of pre-existing desalination plant and the potential risk of the drought ending, our plant will be tentatively designed to produce 15 million gallons of drinkable water per day to homes and businesses in the Los Angeles county area. With a recovery rate of near 30%, and a requirement for diluting any brine, the input feed flowrate will be 70.5 million gallons per day. Following the similarly built Carlsbad plant, we expect the plant to last 30 years. The project building timeline is one year based on our relatively small plant size. We will run for 350 days a year while shutting down for cleaning and maintenance once every 3 months (based on DOW standards).<ref name = "Bates">Bates, Wayne. “Cleaning your RO”. Hydranautics. </ref> The plant will be built 500 meters from the coastline on Long Beach. A borehole will be drilled 10 meters into the ground. At this depth there should be seawater aquifers below the coast that regularly recharges itself with seawater from the ocean. A vertical well containing intake holes will be put into the borehole. The intake holes will be equipped with large size filters that can easily be replaced. At the surface, the well will be connected to a large-scale pump in a well pump house. From there, the seawater will be horizontally pumped into a seawater tank in the desalination plant. This will be considered the start of our process. The process ends after pretreatment, desalination, post treatment, and brine treatment are completed. Appendix 1 shows more details on the Design Basis.

The feed composition is assumed to be the standard seawater “average” based on the DOW Chemical ROSA 9.0 Technical Manual in Table 1. However, in reality the seawater composition at the plant site will have to be chemically tested. The final product will be based on L.A. standards for drinking water quality (Table 1). The specifications in the table are upper limits that we cannot exceed. We will aim to reach the salt level standards but any remineralization and pH treatment will be exported and paid for due to the added complexity and cost in making such a system in-house.

Due to the acidity of the desalinated water coming out of the reverse osmosis (RO) system (pH ~ 6), the pumps and pipes carrying the desalinated water for post treatment will likely corrode over time. We have decided to place the desalination plant by the Long Beach Water Reclamation Plant (about 50 meters away) to prevent pipe corrosion.

[[File:Table_1.jpg|center|200px|frame|Table 1. Concentration of the standard seawater, the output from the RO system, and LA standard.]]

=Desalination Overview and Alternatives=

We first began our research into the current landscape of the desalination industry, particularly of plants that were built in the United States. We discovered there were many processes but the main ones included: 
* Membrane Separation (Forward Osmosis and Reverse Osmosis)
* Thermal Separation (Multi Stage Flash Distillation)
* Phase Change Separation (Freeze Separation)
Each process was analyzed for their pros and cons. For our project design, we decided to go with reverse osmosis as our desalination process. This is because it is the most commonly used process in the industry, and for good reason. Unlike thermal separation, which requires huge amounts of energy to supply heat, reverse osmosis is very energy efficient. This reduces energy costs that would be spread across the plant’s lifetime.<ref name = "Republic">Republic of Mauritius NY. "Technical Aspects of Desalination Plant". Sustainable Sanitation and Water Management. </ref>

Forward osmosis is a newer technology being developed that also uses membranes to separate water and salt and may be more energy efficient overall than reverse osmosis. This is because forward osmosis doesn’t require a pump to pressurize seawater through the membranes, but rather leaves a higher salt concentrated solution (AKA draw solution) on the other side of the membrane for the seawater to be drawn into. Then a second step is required to regenerate the concentrated draw solution and produce purified water.<ref>Mallinson, Alissa. "Study Shows Forward Osmosis Desalination Not Energy Efficient." MIT News [Boston]. July 23, 2014.</ref> We decided to move away from forward osmosis because it is a process that is still being refined and developed for desalination and lacks the data, programs, and technologies that reverse osmosis users benefit from. Additionally, a recent study from MIT has claimed that due to the inherent nature of forward osmosis requiring that seawater be drawn to a higher concentrated salt solution before the water and salt is separated, it cannot possibly be more energy efficient than reverse osmosis.<ref>O'Hern, Sean C., Karnik, Rohit. Nanofiltration across Defect-Sealed Nanoporous Monolayer Graphene. American Chemical Society 15.5 (2014); 3254-260. </ref>

Other processes such as freeze separation were too restrictive in their feed and process demands and required a lot of capital. Ultimately, reverse osmosis is simply the best process that exists for desalination and has been ranked the best when performance is evaluated in terms of technical, environmental, and economic aspects. It is a process that is flexible in water quantity, operation start-up and shut-off, and site location.

Nevertheless there are disadvantages to reverse osmosis that must be considered. The feed seawater must not be too salty and too low quality. The pressure to pump the seawater is enormous. The construction time is long for large plants. The capital costs are also high when compared with other processes such as thermal separation. But all of these disadvantages are ultimately outweighed by the advantages reverse osmosis provides.<ref name = "Republic"/>

=Process Overview and Alternatives=

See Appendix 2 (Figure 1) for the Process Flow Diagram (PFD). The process itself is detailed below. The stream flow rates are in Appendix 2 (Table 1). We assumed 10 meters of pipe for every stream except the inlet and outlet streams of the entire PFD.

==Pretreatment==

After the 10 meter vertical well pumps seawater from the Long Beach coast, the feed is pumped 500 meters to a seawater tank that makes it easy to shut-down the system and reduce the large volume of flow. The level control keeps the seawater from overfilling the tank. From there, part of the seawater goes to brine treatment while the rest go to pretreatment. The purpose of pretreatment in the desalination process is to prevent equipment damage by removing particles and microorganisms from the water. A variety of pollutants -- including calcium carbonate, sulfates, silica, clay, bacteria, and other colloids and microorganisms -- can disrupt the reverse osmosis process. Media filtration uses fine grains such as sand or anthracite to filter the particles. We used a large filter to catch larger particle material (e.g. sand). However, simple filtration is not enough to prevent colloid fouling. An ultrafilter (i.e. microfiltration) is used to remove virtually all suspended matter, effectively preventing colloid fouling. The costs associated with ultrafiltration are not prohibitive in terms of the project’s budget.<ref name = "Desalination">Desalination Pre-treatment. Lenntech.</ref>

Reverse osmosis systems can also be disrupted by the growth of microbiological films of bacteria on the film and pipe walls. They are typically treated as colloids during pretreatment, and most organisms are removed from the water by ultrafiltration, but this does not remove the smallest of organic particles. DBNPA is the most effective biocide, but these aren’t often used and have to be removed from the water. Chlorination is an effective way to remove bacteria, but it can cause degradation of the membrane, and must be removed before the reverse osmosis occurs. In our process we will use UV irradiation to remove any organic material that passes through the microfilters.<ref name = "Desalination"/>

Scaling can reduce the flux across the membrane and can build up in pipe corners, causing backup. It is caused by inorganic salts -- specifically sulfates and calcium and magnesium carbonate -- present in the water. Scaling can be prevented and/or controlled with acids, antiscalants, or ion exchange softening. Antiscalants are good for removing sulfates, but not much else. Acidification will only be effective at removing CaCO3 from water but, the decreased pH may deteriorate the membrane and pipes, so it may require future research.<ref name = "Desalination"/> Ion exchange is very expensive and requires more capital to undergo, but is overall the more effective option.

The ion exchanger is a pretreatment system which acts as a method to fight membrane scaling in reverse osmosis. The ion exchanger works by using sodium carboxylate as a resin which will react with Chlorine and Magnesium ions, removing them from the feed. The sodium is substituted out and enters the feed stream, increasing the overall sodium level and TDS. The lack of magnesium and chlorine ions means that there is nothing to react with the carbonates to form scaling salts on the membranes.

==Desalination (Reverse Osmosis)==

Two pumps (P-105 & P-106) take the adjusted feed seawater up to 58 bar. However, before this pumping occurs, about 70% of the original stream gets separated and heads into a PX Pressure Exchanger. This separated stream later joins together with the pumped stream to enter the reverse osmosis system at the required pressure of 58 bar. The streams can combine together as they have equal pressure. It is the concentrate stream from the output of the reverse osmosis system that transfers its energy to the adjusted feed in the PX Pressure Exchanger.

[[File:Figure_3.jpg|left|300px|frame|Figure 3. Inside look of pressure exchanger]]
The pressure of the concentrate stream is about 57 bar and it is fed into an exchanger which works by using vessel that resembles a small pipe within a rotating wheel. The low pressure (LP) feed enters the vessel at 1 bar (see Figure 3). The feed turns and is pushed out into an outlet pipe by the high pressure (HP) concentrate stream at ~57 bar. The feed is now a HP stream via this exchanged energy. The now LP concentrate stream turns and is pushed out into an outlet pipe by the LP feed entering the vessel.

98% of the energy within the concentrate stream can be recovered via the PX Pressure Exchanger.<ref>Desalination Product and Applications Catalog 2013. “Advance Technologies for Seawater Reverse Osmosis and Brackish Water Applications”. Energy Recovery, Inc. 2012-2014.</ref >Thus, the now HP feed stream is not quite 58 bar and will need to be pumped by a small booster pump (P-107).

When designing a reverse osmosis system, we need to make decisions on what specific kind of reverse osmosis system will be implemented. When making the decision, we investigated the advantages of different types of reverse osmosis among the following options: batch vs. continuous process, number of stages, and number of elements.<ref>FILMTEC Reverse Osmosis Membranes. “Water & Process Solutions Technical Manual.” Dow Chemical Company.</ref> Our decision was made based on process optimization which is described later in this report. The end result of that optimization was to go with a single stage, continuous process with five elements in series.

=Post Treatment=

The permeate is sent to post treatment after the salt content is analyzed and controlled if too high or too low due to any unexpected errors. As stated above, the post treatment process will be done 50 meters from our plant by the Long Beach Water Reclamation Plant. The reason we aren’t performing post treatment “in-house” is due to the increased cost it would take for us to build and run the equipment for post treatment. We have the option to use the government’s water treatment facilities that already exist and our 15 million gallons per day output is a relatively small addition to the large Long Beach Water Reclamation Plant. We will assume any costs to post treatment will be pushed to the consumer in the price of purchasing the drinking water from the Long Beach Water Reclamation Plant.

Post treatment at any facility works to bring water to “taste standards”. The water that is the permeate of our process is already meeting L.A. drinking standards as shown above in Table 1 after our mass balance calculations around the reverse osmosis system (Appendix 3, Table 1). The post treatment increases the pH from 6.3 to 7.3 with the addition of base and then adds minerals like fluoride, calcium, and magnesium back in to improve taste.

==Brine Treatment and Byproduct Waste==

A natural consequence of a reverse osmosis procedure is that a lot of salt will be leftover from the membrane separation in a very concentrated form. This is called “brine” and will have to be taken care of. According to EPA requirements, the salt concentration of water entering the environment (via land or sea) cannot be higher than 40,000 ppm.8 The salt concentration of our brine is much higher (around 70,000 ppm). Thus, the brine must be removed or treated. We decided to treat the brine by diluting it as removing brine can be difficult and may require lots of land (e.g. evaporation ponds).

After the concentrate exchanges energy in the energy exchanger, it becomes a low pressure stream. It is pumped into a dilution tank and mixes with 18.2 million gallons per day of seawater from the inlet to become diluted brine which is pumped 600 meters out to sea.

The waste byproducts are the solid buildup on the filters and membranes. The filters and membranes also have a set lifetime and thus will also need to be removed. The waste is all non-hazardous and thus can be put in the municipal dump.

=Process Details and Equipment Sizing=

==Filters and UV Treatment [Pretreatment]==
The cost of our ultrafiltration system is estimated using data from the University of Tennessee at Knoxville.<ref>Qiang He and R. Bruce Robinson Membrane Filtration Systems. “Membrane Module Costs.” University of Tennessee at Knoxville. December 2014.</ref> It estimated the cost of the cost of the ultrafiltration $0.057 per 1000 gallons and the construction and capital costs are $0.39 per gallon per day. The cost of UV treatment is based off of units available from buyultraviolet.com for units that can treat about 450 gallons per minute.<ref>Atlantic Ultraviolet Corporation. “Megatron Water Disinfection System.”</ref> Due to the water having to be in relatively small pipes for treatment we do not expect the costs to scale very much, so we believe these units to be an accurate cost estimate.

==Ion Exchanger [Pretreatment] ==

The ion exchanger was sized using data from DOW. Our process will use 60 vessels with ion exchange beds of volume 5.73 m3. With a flow rate of 30 bed volumes/hour (10,312.2 m3/hr) the system of beds has a capacity of 125% of the feed. This allows for periodic regeneration of the resin, a weak acid cation resin (DOW MAC-3 Resin). This weak cation resin will remove cationic calcium and magnesium from the water to prevent scaling. The mass flow was calculated via a mass balance on the ion exchanger (Appendix 3).

==Reverse Osmosis System [Desalination]==

Each of the 2607 Reverse Osmosis SW30HRLE-440i [Seawater 30 High Rejection Low Energy 440 ft2 interlocking cap membrane] units have the same dimensions. A, the net length, is 1016 mm. B, the total length including the endcap, is 1029 mm. C, the diameter of the filter element, is 201mm. D, the diameter of the permeate pathway, is 29 mm. For a single stage, we have five of the elements in series so the overall length of a single stage is expected to be roughly 5145 mm. The membrane material is polyamide thin film composite. It will last three years.
Operating Conditions [ROSA 9.0]: 
* Temperature: 30oC 
* Pressure: 58 bar 
* Flow Rate: 52.3 Million gal per day 
* Recovery: 28.7%

[[File:Figure_4.png|center|400px|frame|Figure 4. Diagram of Dow FILMTEC SW30HRLE-440i]]
Cleaning Schedule: 
* Conservative Estimate: Once every three months (three days available for cleaning)<ref name = "Bates"/> 
* DOW FILMTEC Cleaning Conditions 
** The normalized permeate flow drops 10%
* The normalized salt passage increases 5 - 10%

==Seawater Holding Tank and Dilution Tank [Brine Treatment]==

The initial seawater holding tank works as a buffer to prevent surges and help perform shutdowns quickly and easily. We expect to run continuously for 350 days a year with the rest of the available time spread out and used for shutting down the plant for cleaning membranes, filters, and the fixing of unexpected errors. There is a big tradeoff between having a holding tank with high retention time and having a holding tank with low cost. With that in mind, we erred on the side of lowering capital cost and went with a retention time of one hour. The volume could be found by multiplying the retention time with the feed flow rate (Stream 1). We assumed that a good height to width ratio for a large water tank would be 4:1.

A mass balance was done around the dilution tank to find what amount of seawater would be necessary to dilute the brine to 45,000 ppm salinity (Appendix 3; Table 3). With this the feed flow rate into the tank could be found. It was found that it takes about 20 minutes of retention time for good dilution of the brine in seawater.8 We again assumed that a good height to width ratio for a large water tank would be 4:1.

We assumed a corrosion allowance of 4 mm wall thickness. We calculated wall thickness using vessel pressure, hoop stress, longitudinal stress, and weld factor. Each tank is made with SS316 to protect against corrosion.<ref>Towler G, Sinnott R. Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013.</ref>

<center>
'''Table 2.''' Tank sizes.
[[File:Table_2.jpg|center|600px]]
</center>

=Process Optimization=
<center> '''Table 3.'''Optimization Criteria
[[File:Table_3.jpg|center|400px]]
</center>
In order to optimize the reverse osmosis system, numerous simulations were run. Three representative cases are reported as shown above (Table 3). The objective of the optimization was to minimize the required water price to break even in 10 years. In order to achieve this goal, the balance between the capital cost from the number of the filter elements and utility cost from energy consumption was weighed. Before the energy recovery system was implemented, the general trend was that if we have more filter elements in series, overall recovery increases but recovery rate per filter decreases, and we need more filter elements. While increasing the elements in series increase the required number of filters, and thereby increases the capital cost, we require less energy to produce the same amount of water. After the energy recovery system is implemented, the correlation between the number of elements in series and the required energy becomes more complicated because high pressure concentrate is used to pressurize the feed. However, the trade-off between the utility cost and capital cost still exists. Of the three cases shown above, single stage, five elements RO system turned out to be the most optimal solution because single stage, single element system required higher utility cost and two stage system slightly did not meet the LA standard. Detailed information which was used for the optimization is available in Appendix 3 and in Appendix 6.

=Ethical, Environmental, and Societal Considerations=

There are a few key considerations we have followed to meet ethical, environmental, and societal guidelines and requirements: 
* Meet L.A. Drinking Water Standards
** Correct ion concentration
* Moved near water treatment plant
** Pipes will not corrode from low pH
* Working below pressure limits
* Water will be locally served and meet taste standards
** Via Post Treatment mineralization
* Follows EPA Environmental Salinity Requirements
* Reduced Energy Requirements
** PX Pressure Exchanger
** Optimized RO System
** Reverse Osmosis is itself Energy Saving

For further environmental considerations, see effluent summary in Appendix 4.

=Economic Analysis=

The pricing of the majority of the process was completed using the Aspen Economic Analyzer, however we had determine costs for the ultrafiltration, reverse osmosis membranes, and the ultraviolet treatment was completed separately. This led to a total ISBL of $56 MM, OSBL of $22.4 MM, Engineering Costs of $22.4 MM and a Contingency of $7.8 MM. We also chose a working Capital of $10 MM. This working capital is sufficient because we have no raw materials costs. We also had FCOP of $5.2 MM and VCOP of $3.5 MM.

We originally were planning on pricing our water at $700 per acre-foot, but after doing economic analysis we had a 10 year NPV of $-76.2 MM. After determining did a sensitivity analysis and determined that if our water was priced at $2104 per acre-foot we could have an NPV of $0 after 10 years and if we were able to get the same price of $2257 per acre-foot as the plant in Carlsbad we would have a 10 year NPV of $8.3 MM which increases to $41.1 MM at 20 years. For more detailed data about the economic analysis see Appendix 6.

=Conclusion and Recommendations=

Our mission was to create an economically viable process that will desalinate water for human consumption in the L.A. County. We have found after research and optimization that the best process to undergo desalination is a continuous single stage, five elements reverse osmosis system. Our process can meet the required L.A. drinking water standard and we have set up plans for post treatment at a local water reclamation site. While meeting the drinking water standards, we also met our ethical, environmental, and social standards.

Nevertheless, at a sale price of $700 per acre-foot, our proposed desalination plant in Los Angeles, California looks to be a very expensive project to undergo. Without any sort of government assistance to carry the heavy capital costs of such a project, we would suggest the plant not be built. However, after performing a sensitivity analysis at higher (and less conservative) price points we reached a breakeven point in 10 years with a sale price of $2104 per acre-foot of drinking water. Research has shown that the droughts striking California look only to get worse in the near future. With limited freshwater resources, there is no option other than to consider desalination. Therefore it is reasonable to expect we could get near the same price deal as local desalination plants like the Carlsbad Plant ($2257 per acre-foot) As we have met all design constraints for this project, we hope to convince policymakers to adopt our project at least $2104 per acre-foot water. If this deal is reached, we recommend that this project and design be funded.

However, there are ways to improve our design with future work. Right now all the pipes diameters are based on having a flowrate of 1.5 m/s. But in reality, we’d need to use standard piping sizes to get more accurate pricing and energy costs. Also, in reality we would be purchasing some of these vessels in bulk quantities (i.e. bulk filters) so the price could be lowered with negotiation and economies of scale. We also have pumps that are providing very little pressure for their high cost. We can remove these pumps and simply run at higher pressures earlier. We can potentially reduce early energy costs by building a “floating” desalination plant in the sea to reduce pumping distance from the well.

=Appendices=

==Appendix 1: Design Basis==
[[File:Appendix_1.png|center|Appendix_1.png]]

==Appendix 2: Process Flow Diagram==
[[File:Appendix_2.png|center]]
==Appendix 3: Mass and Energy Balance Tables==
'''Table 1'''. Mass balance on 1 stage 5 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T1.png|center|500px|]]

'''Table 2'''. Mass balance on 1 stage 1 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T2.png|center|500px|]]

'''Table 3'''. Mass balance on 2 stage 5,5 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T3.png|center|600px|]]

'''Table 4'''. 1 stage 5 elements CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T4.png|center|300px|]]

'''Table 5'''. 1 stage 1 element CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T5.png|center|300px|]]

'''Table 6'''. 2 stages 5,5 elements CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T6.png|center|300px|]]

'''Table 7'''. Energy balance on 1 stage 5 elements RO system
[[File:Appendix_3_T7.png|center|800px|]]

==Appendix 4: Effluent Summary==
[[File:Appendix_4.png|center]]

==Appendix 5: Cost of Production==
[[File:Appendix_5.png|center]]

==Appendix 6: Economic Anlaysis==
===Case I ===
Water is $700; RO with 1 stage and 5 elements
[[File:Appendix_6_C1.png|center]]

===Case II===
Water is $1400; RO with 1 stage and 5 elements
[[File:Appendix_6_C2.png|center]]

===Case III: Carlsbad ===
Water is $2257; RO with 1 stage and 5 elements
[[File:Appendix_6_C3.png|center]]

===Case IV: Breakeven (NPV = 0 at 10 years)===
Water is $2104; RO with 1 stage and 5 elements
[[File:Appendix_6_C4.png|center]]

===Case V: Breakeven (NPV = 0 at 10 years)===
Water is $2112; RO with 2 stages and 5,5 elements
[[File:Appendix_6_C5.png|center]]

===Case VI: Breakeven (NPV = 0 at 10 years)===
Water is $2138 with 1 stage and 1 element
[[File:Appendix_6_C6.png|center]]

==Appendix 7: ROSA Simulation Outputs==

'''Table 1'''. System detail of 1 stage 5 elements RO system
[[File:Appendix_7_T1.png|center]]

'''Table 2'''. System detail of 1 stage 1 elements RO system
[[File:Appendix_7_T2.png|center]]

'''Table 3'''. System detail of 2 stage 5,5 elements system
[[File:Appendix_7_T3.png|center]]

'''Table 4'''. 1 stage 5 elements RO system concentrations and pH in each stream
[[File:Appendix_7_T4.png|center]]

'''Table 5'''. 1 stage 1 element RO system concentration and pH in each stream
[[File:Appendix_7_T5.png|center]]

'''Table 6'''. 2 stage 5,5 elements RO system concentrations and pH in each stream
[[File:Appendix_7_T6.png|center]]

=References=
<references/>

Desalination - Team E

2016-03-12T03:35:54Z

Bts511: /* Appendix 1: Design Basis */

Group E Corporation

Authors: Hassan Ali, Woo Soo Choe, Brett Sleyster, Jake Stolley

Instructors: Fengqi You, David Wegerer

March 11, 2016

=Executive Summary=

Outlined in this report is the project model report for the proposed desalination plant in Los Angeles, California which will provide drinking water to local citizens using reverse osmosis process. This location was chosen due to the region’s propensity for droughts in recent years. The fresh water resources in California have plummeted as water prices have risen 3-4% each year. This is only expected to continue. Other companies have worked deals with the Californian government to create desalination plants that provide a new fresh water drinking source. This suggests that there is a strong market here with potential growth.

The plant will have an output of 15 million gallons per day, which is about medium size relative to other plants. Thus the build time will be about one year. The plant will run for 30 years, with 350 days per year of use. Shutdown time will be used for cleaning and maintenance of equipment. The plant will be built on Long Beach, 500 meters from the coast nearby the Long Beach Water Reclamation Plant (LBWP). The input will withdraw 70.5 million gallons per day of seawater that will be used for reverse osmosis and dilution of brine.

Reverse osmosis was chosen as the desalination method for its low energy cost and flexibility. We designed a process that will pretreat the seawater to remove waste, colloids, bacteria, and membrane-damaging scalants. Once this seawater is treated is pumped to heavy pressures through a reverse osmosis system. After doing process optimization under various system parameter changes, we found a continuous, single stage, five element reverse osmosis process to be best. It produces a near 30% recovery. We reduce energy costs by including a pressure exchanger between our concentrate outlet and adjusted feed into the reverse osmosis system.

After performing mass and energy balances, we were able to size equipment and find our permeate composition. It turns out our process will meet L.A. standards for drinking water quality. However, post treatment will be done by the local LBWP to make the water taste better and be non-caustic to pipes. This will reduce “in-house” costs. Many environmental considerations were made, to meet our environmental standards. We meet EPA salinity requirements for our brine waste by diluting it. We also don’t use a lot of energy and thus reduce our fossil fuel consumption and CO2 generation. To be safe, we work well under max pressure limits for our equipment. We make sure we are ethical and meet all standards that our local county requires of us.

After an economic evaluation at a conservative well-water price of $700 per acre-foot, we realize that our revenue is way too low to cover our costs. Capital costs are especially large with regards to our large water tanks. We find that we reach a 10 year NPV breakpoint of $2104 per acre-foot after performing a sensitivity analysis. The Carlsbad Plant in San Diego worked out a deal for $2257 per acre-foot.

As we have met all design constraints for this project, we hope to convince policymakers to adopt our project at least $2104 per acre-foot water. If this deal is reached, we recommend that this project and design be funded. There is future work to be done to improve costs such as performing pump optimization. However, this work is minimal and won’t harm our bottom line if we reach our proposed price points.

__TOC__

=Introduction=

Due to a combination of various factors--including rising human population, climate change, increased energy consumption, and land erosion--the availability of fresh water as a resource for human consumption has been dwindling. This has had a negative impact on the world, but this impact has been especially felt in areas with natural shortages in fresh water. Desert regions, such as Saudi Arabia, have resorted to desalinating seawater at the coast and transporting up to their cities. This encompasses their entire freshwater system as very few natural water sources remain. In the United States there is a region which has shown increasing signs of following in the footsteps of Saudi Arabia and becoming fully reliant on desalination: Los Angeles, California.

=Location and Market Analysis=

In this project, Los Angeles was chosen as the location our plant, because the area has been subject to a severe drought for the last four years, as shown in Figure 1. The rapid decrease in the available water has caused permanent damages to the ground, decreased the amount of water in the reservoir from 45% to 25%, and reduced the amount of hydroelectric power produced in the state. Such deficiency in usable water has increased the water prices in Los Angeles by 6-7% annually for the last decade and is expected to rise by 3.4% annually for typical customers for the next five years. For customers who use more than 20000 gallons a year, the bill is expected to increase by 34% by 2021. This price increase is occurring because of both drought and maintenance. <ref>Pacific Institute: The California Drought. 2014.</ref>

[[File:Figure_1.jpg|right|200px|frame|Figure 1. California Drought Index.]]
As of March 2015 the Metropolitan Water District of Southern California was willing to pay as much as $700 per acre-foot of water to farmers in the Sacramento Valley, so we assume this will be the target price of the water produced by the this desalination plant.<ref>Hessel, Phil. “Dry Southern California Offers Northern Farmers Top Dollar for Water”. NBC News. March 18, 2015. </ref>

Currently, there is a plant in Carlsbad which supplies drinking water to the San Diego area. The plant can produce up to 50 million gallons per day and is the nation's largest water desalination plant. The water from this plant costs $2257 per acre-foot for the first 48000 acre-feet and about $2000 for any additional acre-foot. The county signed a contract with this plant to purchase the first 48000 acre-feet of water for the next 30 years.<ref>Carlsbad Desalination Plant. “Project Overview”. January 2016.</ref> Also, there is another desalination plant in Tampa that produces about 25 million gallons of water per day, and over 2000 other public and industrial plants with production capacity of over 300000 gallons per day. When conducting the economic analysis our plant, such numbers presented by other plants must be considered, so our project can yield positive profit.<ref>Leven, Rachel. “U.S. Desalination Industry Grows Since 2000; See as Essential to Meeting Supply Needs”. Bloomberg BNA. August 21, 2013.</ref>

The profitability of the project is naturally heavily dependent on the price of water in California. If the drought ends, water supply will increase and the water price will drop significantly, but if the drought continues, the price of water in California could rise even more rapidly and allow desalination plants to be more profitable.<ref>Southern California Public Radio. “Where is California water use decreasing?”. December 2015.</ref> It may be difficult to make this plant profitable without a guaranteed purchase or other subsidy from the local governments because the presence of pre-existing plants maintains the water price relatively low. However, external sources of profit such as guaranteed purchase or subsidy are expected because the ramifications of not having desalination plant or other sources of water may be significantly more expensive than the subsidizing a desalination plant assuming the drought continues.<ref>Weiser, Matt. “Could desalination solve California’s water problem?”. The Sacramento Bee. October 18, 2014.</ref>

=Design Basis and Technical Approach=

[[File:Figure_2.jpg|left|200px|frame|Figure 2. Vertical Well Feed]]
Considering the number of pre-existing desalination plant and the potential risk of the drought ending, our plant will be tentatively designed to produce 15 million gallons of drinkable water per day to homes and businesses in the Los Angeles county area. With a recovery rate of near 30%, and a requirement for diluting any brine, the input feed flowrate will be 70.5 million gallons per day. Following the similarly built Carlsbad plant, we expect the plant to last 30 years. The project building timeline is one year based on our relatively small plant size. We will run for 350 days a year while shutting down for cleaning and maintenance once every 3 months (based on DOW standards).<ref name = "Bates">Bates, Wayne. “Cleaning your RO”. Hydranautics. </ref> The plant will be built 500 meters from the coastline on Long Beach. A borehole will be drilled 10 meters into the ground. At this depth there should be seawater aquifers below the coast that regularly recharges itself with seawater from the ocean. A vertical well containing intake holes will be put into the borehole. The intake holes will be equipped with large size filters that can easily be replaced. At the surface, the well will be connected to a large-scale pump in a well pump house. From there, the seawater will be horizontally pumped into a seawater tank in the desalination plant. This will be considered the start of our process. The process ends after pretreatment, desalination, post treatment, and brine treatment are completed. Appendix 1 shows more details on the Design Basis.

The feed composition is assumed to be the standard seawater “average” based on the DOW Chemical ROSA 9.0 Technical Manual in Table 1. However, in reality the seawater composition at the plant site will have to be chemically tested. The final product will be based on L.A. standards for drinking water quality (Table 1). The specifications in the table are upper limits that we cannot exceed. We will aim to reach the salt level standards but any remineralization and pH treatment will be exported and paid for due to the added complexity and cost in making such a system in-house.

Due to the acidity of the desalinated water coming out of the reverse osmosis (RO) system (pH ~ 6), the pumps and pipes carrying the desalinated water for post treatment will likely corrode over time. We have decided to place the desalination plant by the Long Beach Water Reclamation Plant (about 50 meters away) to prevent pipe corrosion.

[[File:Table_1.jpg|center|200px|frame|Table 1. Concentration of the standard seawater, the output from the RO system, and LA standard.]]

=Desalination Overview and Alternatives=

We first began our research into the current landscape of the desalination industry, particularly of plants that were built in the United States. We discovered there were many processes but the main ones included: 
* Membrane Separation (Forward Osmosis and Reverse Osmosis)
* Thermal Separation (Multi Stage Flash Distillation)
* Phase Change Separation (Freeze Separation)
Each process was analyzed for their pros and cons. For our project design, we decided to go with reverse osmosis as our desalination process. This is because it is the most commonly used process in the industry, and for good reason. Unlike thermal separation, which requires huge amounts of energy to supply heat, reverse osmosis is very energy efficient. This reduces energy costs that would be spread across the plant’s lifetime.<ref name = "Republic">Republic of Mauritius NY. "Technical Aspects of Desalination Plant". Sustainable Sanitation and Water Management. </ref>

Forward osmosis is a newer technology being developed that also uses membranes to separate water and salt and may be more energy efficient overall than reverse osmosis. This is because forward osmosis doesn’t require a pump to pressurize seawater through the membranes, but rather leaves a higher salt concentrated solution (AKA draw solution) on the other side of the membrane for the seawater to be drawn into. Then a second step is required to regenerate the concentrated draw solution and produce purified water.<ref>Mallinson, Alissa. "Study Shows Forward Osmosis Desalination Not Energy Efficient." MIT News [Boston]. July 23, 2014.</ref> We decided to move away from forward osmosis because it is a process that is still being refined and developed for desalination and lacks the data, programs, and technologies that reverse osmosis users benefit from. Additionally, a recent study from MIT has claimed that due to the inherent nature of forward osmosis requiring that seawater be drawn to a higher concentrated salt solution before the water and salt is separated, it cannot possibly be more energy efficient than reverse osmosis.<ref>O'Hern, Sean C., Karnik, Rohit. Nanofiltration across Defect-Sealed Nanoporous Monolayer Graphene. American Chemical Society 15.5 (2014); 3254-260. </ref>

Other processes such as freeze separation were too restrictive in their feed and process demands and required a lot of capital. Ultimately, reverse osmosis is simply the best process that exists for desalination and has been ranked the best when performance is evaluated in terms of technical, environmental, and economic aspects. It is a process that is flexible in water quantity, operation start-up and shut-off, and site location.

Nevertheless there are disadvantages to reverse osmosis that must be considered. The feed seawater must not be too salty and too low quality. The pressure to pump the seawater is enormous. The construction time is long for large plants. The capital costs are also high when compared with other processes such as thermal separation. But all of these disadvantages are ultimately outweighed by the advantages reverse osmosis provides.<ref name = "Republic"/>

=Process Overview and Alternatives=

See Appendix 2 (Figure 1) for the Process Flow Diagram (PFD). The process itself is detailed below. The stream flow rates are in Appendix 2 (Table 1). We assumed 10 meters of pipe for every stream except the inlet and outlet streams of the entire PFD.

==Pretreatment==

After the 10 meter vertical well pumps seawater from the Long Beach coast, the feed is pumped 500 meters to a seawater tank that makes it easy to shut-down the system and reduce the large volume of flow. The level control keeps the seawater from overfilling the tank. From there, part of the seawater goes to brine treatment while the rest go to pretreatment. The purpose of pretreatment in the desalination process is to prevent equipment damage by removing particles and microorganisms from the water. A variety of pollutants -- including calcium carbonate, sulfates, silica, clay, bacteria, and other colloids and microorganisms -- can disrupt the reverse osmosis process. Media filtration uses fine grains such as sand or anthracite to filter the particles. We used a large filter to catch larger particle material (e.g. sand). However, simple filtration is not enough to prevent colloid fouling. An ultrafilter (i.e. microfiltration) is used to remove virtually all suspended matter, effectively preventing colloid fouling. The costs associated with ultrafiltration are not prohibitive in terms of the project’s budget.<ref name = "Desalination">Desalination Pre-treatment. Lenntech.</ref>

Reverse osmosis systems can also be disrupted by the growth of microbiological films of bacteria on the film and pipe walls. They are typically treated as colloids during pretreatment, and most organisms are removed from the water by ultrafiltration, but this does not remove the smallest of organic particles. DBNPA is the most effective biocide, but these aren’t often used and have to be removed from the water. Chlorination is an effective way to remove bacteria, but it can cause degradation of the membrane, and must be removed before the reverse osmosis occurs. In our process we will use UV irradiation to remove any organic material that passes through the microfilters.<ref name = "Desalination"/>

Scaling can reduce the flux across the membrane and can build up in pipe corners, causing backup. It is caused by inorganic salts -- specifically sulfates and calcium and magnesium carbonate -- present in the water. Scaling can be prevented and/or controlled with acids, antiscalants, or ion exchange softening. Antiscalants are good for removing sulfates, but not much else. Acidification will only be effective at removing CaCO3 from water but, the decreased pH may deteriorate the membrane and pipes, so it may require future research.<ref name = "Desalination"/> Ion exchange is very expensive and requires more capital to undergo, but is overall the more effective option.

The ion exchanger is a pretreatment system which acts as a method to fight membrane scaling in reverse osmosis. The ion exchanger works by using sodium carboxylate as a resin which will react with Chlorine and Magnesium ions, removing them from the feed. The sodium is substituted out and enters the feed stream, increasing the overall sodium level and TDS. The lack of magnesium and chlorine ions means that there is nothing to react with the carbonates to form scaling salts on the membranes.

==Desalination (Reverse Osmosis)==

Two pumps (P-105 & P-106) take the adjusted feed seawater up to 58 bar. However, before this pumping occurs, about 70% of the original stream gets separated and heads into a PX Pressure Exchanger. This separated stream later joins together with the pumped stream to enter the reverse osmosis system at the required pressure of 58 bar. The streams can combine together as they have equal pressure. It is the concentrate stream from the output of the reverse osmosis system that transfers its energy to the adjusted feed in the PX Pressure Exchanger.

[[File:Figure_3.jpg|left|300px|frame|Figure 3. Inside look of pressure exchanger]]
The pressure of the concentrate stream is about 57 bar and it is fed into an exchanger which works by using vessel that resembles a small pipe within a rotating wheel. The low pressure (LP) feed enters the vessel at 1 bar (see Figure 3). The feed turns and is pushed out into an outlet pipe by the high pressure (HP) concentrate stream at ~57 bar. The feed is now a HP stream via this exchanged energy. The now LP concentrate stream turns and is pushed out into an outlet pipe by the LP feed entering the vessel.

98% of the energy within the concentrate stream can be recovered via the PX Pressure Exchanger.<ref>Desalination Product and Applications Catalog 2013. “Advance Technologies for Seawater Reverse Osmosis and Brackish Water Applications”. Energy Recovery, Inc. 2012-2014.</ref >Thus, the now HP feed stream is not quite 58 bar and will need to be pumped by a small booster pump (P-107).

When designing a reverse osmosis system, we need to make decisions on what specific kind of reverse osmosis system will be implemented. When making the decision, we investigated the advantages of different types of reverse osmosis among the following options: batch vs. continuous process, number of stages, and number of elements.<ref>FILMTEC Reverse Osmosis Membranes. “Water & Process Solutions Technical Manual.” Dow Chemical Company.</ref> Our decision was made based on process optimization which is described later in this report. The end result of that optimization was to go with a single stage, continuous process with five elements in series.

=Post Treatment=

The permeate is sent to post treatment after the salt content is analyzed and controlled if too high or too low due to any unexpected errors. As stated above, the post treatment process will be done 50 meters from our plant by the Long Beach Water Reclamation Plant. The reason we aren’t performing post treatment “in-house” is due to the increased cost it would take for us to build and run the equipment for post treatment. We have the option to use the government’s water treatment facilities that already exist and our 15 million gallons per day output is a relatively small addition to the large Long Beach Water Reclamation Plant. We will assume any costs to post treatment will be pushed to the consumer in the price of purchasing the drinking water from the Long Beach Water Reclamation Plant.

Post treatment at any facility works to bring water to “taste standards”. The water that is the permeate of our process is already meeting L.A. drinking standards as shown above in Table 1 after our mass balance calculations around the reverse osmosis system (Appendix 3, Table 1). The post treatment increases the pH from 6.3 to 7.3 with the addition of base and then adds minerals like fluoride, calcium, and magnesium back in to improve taste.

==Brine Treatment and Byproduct Waste==

A natural consequence of a reverse osmosis procedure is that a lot of salt will be leftover from the membrane separation in a very concentrated form. This is called “brine” and will have to be taken care of. According to EPA requirements, the salt concentration of water entering the environment (via land or sea) cannot be higher than 40,000 ppm.8 The salt concentration of our brine is much higher (around 70,000 ppm). Thus, the brine must be removed or treated. We decided to treat the brine by diluting it as removing brine can be difficult and may require lots of land (e.g. evaporation ponds).

After the concentrate exchanges energy in the energy exchanger, it becomes a low pressure stream. It is pumped into a dilution tank and mixes with 18.2 million gallons per day of seawater from the inlet to become diluted brine which is pumped 600 meters out to sea.

The waste byproducts are the solid buildup on the filters and membranes. The filters and membranes also have a set lifetime and thus will also need to be removed. The waste is all non-hazardous and thus can be put in the municipal dump.

=Process Details and Equipment Sizing=

==Filters and UV Treatment [Pretreatment]==
The cost of our ultrafiltration system is estimated using data from the University of Tennessee at Knoxville.<ref>Qiang He and R. Bruce Robinson Membrane Filtration Systems. “Membrane Module Costs.” University of Tennessee at Knoxville. December 2014.</ref> It estimated the cost of the cost of the ultrafiltration $0.057 per 1000 gallons and the construction and capital costs are $0.39 per gallon per day. The cost of UV treatment is based off of units available from buyultraviolet.com for units that can treat about 450 gallons per minute.<ref>Atlantic Ultraviolet Corporation. “Megatron Water Disinfection System.”</ref> Due to the water having to be in relatively small pipes for treatment we do not expect the costs to scale very much, so we believe these units to be an accurate cost estimate.

==Ion Exchanger [Pretreatment] ==

The ion exchanger was sized using data from DOW. Our process will use 60 vessels with ion exchange beds of volume 5.73 m3. With a flow rate of 30 bed volumes/hour (10,312.2 m3/hr) the system of beds has a capacity of 125% of the feed. This allows for periodic regeneration of the resin, a weak acid cation resin (DOW MAC-3 Resin). This weak cation resin will remove cationic calcium and magnesium from the water to prevent scaling. The mass flow was calculated via a mass balance on the ion exchanger (Appendix 3).

==Reverse Osmosis System [Desalination]==

Each of the 2607 Reverse Osmosis SW30HRLE-440i [Seawater 30 High Rejection Low Energy 440 ft2 interlocking cap membrane] units have the same dimensions. A, the net length, is 1016 mm. B, the total length including the endcap, is 1029 mm. C, the diameter of the filter element, is 201mm. D, the diameter of the permeate pathway, is 29 mm. For a single stage, we have five of the elements in series so the overall length of a single stage is expected to be roughly 5145 mm. The membrane material is polyamide thin film composite. It will last three years.
Operating Conditions [ROSA 9.0]: 
* Temperature: 30oC 
* Pressure: 58 bar 
* Flow Rate: 52.3 Million gal per day 
* Recovery: 28.7%

[[File:Figure_4.png|center|400px|frame|Figure 4. Diagram of Dow FILMTEC SW30HRLE-440i]]
Cleaning Schedule: 
* Conservative Estimate: Once every three months (three days available for cleaning)<ref name = "Bates"/> 
* DOW FILMTEC Cleaning Conditions 
** The normalized permeate flow drops 10%
* The normalized salt passage increases 5 - 10%

==Seawater Holding Tank and Dilution Tank [Brine Treatment]==

The initial seawater holding tank works as a buffer to prevent surges and help perform shutdowns quickly and easily. We expect to run continuously for 350 days a year with the rest of the available time spread out and used for shutting down the plant for cleaning membranes, filters, and the fixing of unexpected errors. There is a big tradeoff between having a holding tank with high retention time and having a holding tank with low cost. With that in mind, we erred on the side of lowering capital cost and went with a retention time of one hour. The volume could be found by multiplying the retention time with the feed flow rate (Stream 1). We assumed that a good height to width ratio for a large water tank would be 4:1.

A mass balance was done around the dilution tank to find what amount of seawater would be necessary to dilute the brine to 45,000 ppm salinity (Appendix 3; Table 3). With this the feed flow rate into the tank could be found. It was found that it takes about 20 minutes of retention time for good dilution of the brine in seawater.8 We again assumed that a good height to width ratio for a large water tank would be 4:1.

We assumed a corrosion allowance of 4 mm wall thickness. We calculated wall thickness using vessel pressure, hoop stress, longitudinal stress, and weld factor. Each tank is made with SS316 to protect against corrosion.<ref>Towler G, Sinnott R. Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013.</ref>

<center>
'''Table 2.''' Tank sizes.
[[File:Table_2.jpg|center|600px]]
</center>

=Process Optimization=
<center> '''Table 3.'''Optimization Criteria
[[File:Table_3.jpg|center|400px]]
</center>
In order to optimize the reverse osmosis system, numerous simulations were run. Three representative cases are reported as shown above (Table 3). The objective of the optimization was to minimize the required water price to break even in 10 years. In order to achieve this goal, the balance between the capital cost from the number of the filter elements and utility cost from energy consumption was weighed. Before the energy recovery system was implemented, the general trend was that if we have more filter elements in series, overall recovery increases but recovery rate per filter decreases, and we need more filter elements. While increasing the elements in series increase the required number of filters, and thereby increases the capital cost, we require less energy to produce the same amount of water. After the energy recovery system is implemented, the correlation between the number of elements in series and the required energy becomes more complicated because high pressure concentrate is used to pressurize the feed. However, the trade-off between the utility cost and capital cost still exists. Of the three cases shown above, single stage, five elements RO system turned out to be the most optimal solution because single stage, single element system required higher utility cost and two stage system slightly did not meet the LA standard. Detailed information which was used for the optimization is available in Appendix 3 and in Appendix 6.

=Ethical, Environmental, and Societal Considerations=

There are a few key considerations we have followed to meet ethical, environmental, and societal guidelines and requirements: 
* Meet L.A. Drinking Water Standards
** Correct ion concentration
* Moved near water treatment plant
** Pipes will not corrode from low pH
* Working below pressure limits
* Water will be locally served and meet taste standards
** Via Post Treatment mineralization
* Follows EPA Environmental Salinity Requirements
* Reduced Energy Requirements
** PX Pressure Exchanger
** Optimized RO System
** Reverse Osmosis is itself Energy Saving

For further environmental considerations, see effluent summary in Appendix 4.

=Economic Analysis=

The pricing of the majority of the process was completed using the Aspen Economic Analyzer, however we had determine costs for the ultrafiltration, reverse osmosis membranes, and the ultraviolet treatment was completed separately. This led to a total ISBL of $56 MM, OSBL of $22.4 MM, Engineering Costs of $22.4 MM and a Contingency of $7.8 MM. We also chose a working Capital of $10 MM. This working capital is sufficient because we have no raw materials costs. We also had FCOP of $5.2 MM and VCOP of $3.5 MM.

We originally were planning on pricing our water at $700 per acre-foot, but after doing economic analysis we had a 10 year NPV of $-76.2 MM. After determining did a sensitivity analysis and determined that if our water was priced at $2104 per acre-foot we could have an NPV of $0 after 10 years and if we were able to get the same price of $2257 per acre-foot as the plant in Carlsbad we would have a 10 year NPV of $8.3 MM which increases to $41.1 MM at 20 years. For more detailed data about the economic analysis see Appendix 6.

=Conclusion and Recommendations=

Our mission was to create an economically viable process that will desalinate water for human consumption in the L.A. County. We have found after research and optimization that the best process to undergo desalination is a continuous single stage, five elements reverse osmosis system. Our process can meet the required L.A. drinking water standard and we have set up plans for post treatment at a local water reclamation site. While meeting the drinking water standards, we also met our ethical, environmental, and social standards.

Nevertheless, at a sale price of $700 per acre-foot, our proposed desalination plant in Los Angeles, California looks to be a very expensive project to undergo. Without any sort of government assistance to carry the heavy capital costs of such a project, we would suggest the plant not be built. However, after performing a sensitivity analysis at higher (and less conservative) price points we reached a breakeven point in 10 years with a sale price of $2104 per acre-foot of drinking water. Research has shown that the droughts striking California look only to get worse in the near future. With limited freshwater resources, there is no option other than to consider desalination. Therefore it is reasonable to expect we could get near the same price deal as local desalination plants like the Carlsbad Plant ($2257 per acre-foot) As we have met all design constraints for this project, we hope to convince policymakers to adopt our project at least $2104 per acre-foot water. If this deal is reached, we recommend that this project and design be funded.

However, there are ways to improve our design with future work. Right now all the pipes diameters are based on having a flowrate of 1.5 m/s. But in reality, we’d need to use standard piping sizes to get more accurate pricing and energy costs. Also, in reality we would be purchasing some of these vessels in bulk quantities (i.e. bulk filters) so the price could be lowered with negotiation and economies of scale. We also have pumps that are providing very little pressure for their high cost. We can remove these pumps and simply run at higher pressures earlier. We can potentially reduce early energy costs by building a “floating” desalination plant in the sea to reduce pumping distance from the well.

=Appendices=

==Appendix 1: Design Basis==
[[File:Appendix_1.png|Appendix_1.png]]

==Appendix 2: Process Flow Diagram==
[[File:Appendix_2.png|center]]
==Appendix 3: Mass and Energy Balance Tables==
'''Table 1'''. Mass balance on 1 stage 5 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T1.png|center|500px|]]

'''Table 2'''. Mass balance on 1 stage 1 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T2.png|center|500px|]]

'''Table 3'''. Mass balance on 2 stage 5,5 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T3.png|center|600px|]]

'''Table 4'''. 1 stage 5 elements CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T4.png|center|300px|]]

'''Table 5'''. 1 stage 1 element CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T5.png|center|300px|]]

'''Table 6'''. 2 stages 5,5 elements CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T6.png|center|300px|]]

'''Table 7'''. Energy balance on 1 stage 5 elements RO system
[[File:Appendix_3_T7.png|center|800px|]]

==Appendix 4: Effluent Summary==
[[File:Appendix_4.png|center]]

==Appendix 5: Cost of Production==
[[File:Appendix_5.png|center]]

==Appendix 6: Economic Anlaysis==
===Case I ===
Water is $700; RO with 1 stage and 5 elements
[[File:Appendix_6_C1.png|center]]

===Case II===
Water is $1400; RO with 1 stage and 5 elements
[[File:Appendix_6_C2.png|center]]

===Case III: Carlsbad ===
Water is $2257; RO with 1 stage and 5 elements
[[File:Appendix_6_C3.png|center]]

===Case IV: Breakeven (NPV = 0 at 10 years)===
Water is $2104; RO with 1 stage and 5 elements
[[File:Appendix_6_C4.png|center]]

===Case V: Breakeven (NPV = 0 at 10 years)===
Water is $2112; RO with 2 stages and 5,5 elements
[[File:Appendix_6_C5.png|center]]

===Case VI: Breakeven (NPV = 0 at 10 years)===
Water is $2138 with 1 stage and 1 element
[[File:Appendix_6_C6.png|center]]

==Appendix 7: ROSA Simulation Outputs==

'''Table 1'''. System detail of 1 stage 5 elements RO system
[[File:Appendix_7_T1.png|center]]

'''Table 2'''. System detail of 1 stage 1 elements RO system
[[File:Appendix_7_T2.png|center]]

'''Table 3'''. System detail of 2 stage 5,5 elements system
[[File:Appendix_7_T3.png|center]]

'''Table 4'''. 1 stage 5 elements RO system concentrations and pH in each stream
[[File:Appendix_7_T4.png|center]]

'''Table 5'''. 1 stage 1 element RO system concentration and pH in each stream
[[File:Appendix_7_T5.png|center]]

'''Table 6'''. 2 stage 5,5 elements RO system concentrations and pH in each stream
[[File:Appendix_7_T6.png|center]]

=References=
<references/>

Desalination - Team E

2016-03-12T03:35:22Z

Bts511: /* Appendix 3: Mass and Energy Balance Tables */

Group E Corporation

Authors: Hassan Ali, Woo Soo Choe, Brett Sleyster, Jake Stolley

Instructors: Fengqi You, David Wegerer

March 11, 2016

=Executive Summary=

Outlined in this report is the project model report for the proposed desalination plant in Los Angeles, California which will provide drinking water to local citizens using reverse osmosis process. This location was chosen due to the region’s propensity for droughts in recent years. The fresh water resources in California have plummeted as water prices have risen 3-4% each year. This is only expected to continue. Other companies have worked deals with the Californian government to create desalination plants that provide a new fresh water drinking source. This suggests that there is a strong market here with potential growth.

The plant will have an output of 15 million gallons per day, which is about medium size relative to other plants. Thus the build time will be about one year. The plant will run for 30 years, with 350 days per year of use. Shutdown time will be used for cleaning and maintenance of equipment. The plant will be built on Long Beach, 500 meters from the coast nearby the Long Beach Water Reclamation Plant (LBWP). The input will withdraw 70.5 million gallons per day of seawater that will be used for reverse osmosis and dilution of brine.

Reverse osmosis was chosen as the desalination method for its low energy cost and flexibility. We designed a process that will pretreat the seawater to remove waste, colloids, bacteria, and membrane-damaging scalants. Once this seawater is treated is pumped to heavy pressures through a reverse osmosis system. After doing process optimization under various system parameter changes, we found a continuous, single stage, five element reverse osmosis process to be best. It produces a near 30% recovery. We reduce energy costs by including a pressure exchanger between our concentrate outlet and adjusted feed into the reverse osmosis system.

After performing mass and energy balances, we were able to size equipment and find our permeate composition. It turns out our process will meet L.A. standards for drinking water quality. However, post treatment will be done by the local LBWP to make the water taste better and be non-caustic to pipes. This will reduce “in-house” costs. Many environmental considerations were made, to meet our environmental standards. We meet EPA salinity requirements for our brine waste by diluting it. We also don’t use a lot of energy and thus reduce our fossil fuel consumption and CO2 generation. To be safe, we work well under max pressure limits for our equipment. We make sure we are ethical and meet all standards that our local county requires of us.

After an economic evaluation at a conservative well-water price of $700 per acre-foot, we realize that our revenue is way too low to cover our costs. Capital costs are especially large with regards to our large water tanks. We find that we reach a 10 year NPV breakpoint of $2104 per acre-foot after performing a sensitivity analysis. The Carlsbad Plant in San Diego worked out a deal for $2257 per acre-foot.

As we have met all design constraints for this project, we hope to convince policymakers to adopt our project at least $2104 per acre-foot water. If this deal is reached, we recommend that this project and design be funded. There is future work to be done to improve costs such as performing pump optimization. However, this work is minimal and won’t harm our bottom line if we reach our proposed price points.

__TOC__

=Introduction=

Due to a combination of various factors--including rising human population, climate change, increased energy consumption, and land erosion--the availability of fresh water as a resource for human consumption has been dwindling. This has had a negative impact on the world, but this impact has been especially felt in areas with natural shortages in fresh water. Desert regions, such as Saudi Arabia, have resorted to desalinating seawater at the coast and transporting up to their cities. This encompasses their entire freshwater system as very few natural water sources remain. In the United States there is a region which has shown increasing signs of following in the footsteps of Saudi Arabia and becoming fully reliant on desalination: Los Angeles, California.

=Location and Market Analysis=

In this project, Los Angeles was chosen as the location our plant, because the area has been subject to a severe drought for the last four years, as shown in Figure 1. The rapid decrease in the available water has caused permanent damages to the ground, decreased the amount of water in the reservoir from 45% to 25%, and reduced the amount of hydroelectric power produced in the state. Such deficiency in usable water has increased the water prices in Los Angeles by 6-7% annually for the last decade and is expected to rise by 3.4% annually for typical customers for the next five years. For customers who use more than 20000 gallons a year, the bill is expected to increase by 34% by 2021. This price increase is occurring because of both drought and maintenance. <ref>Pacific Institute: The California Drought. 2014.</ref>

[[File:Figure_1.jpg|right|200px|frame|Figure 1. California Drought Index.]]
As of March 2015 the Metropolitan Water District of Southern California was willing to pay as much as $700 per acre-foot of water to farmers in the Sacramento Valley, so we assume this will be the target price of the water produced by the this desalination plant.<ref>Hessel, Phil. “Dry Southern California Offers Northern Farmers Top Dollar for Water”. NBC News. March 18, 2015. </ref>

Currently, there is a plant in Carlsbad which supplies drinking water to the San Diego area. The plant can produce up to 50 million gallons per day and is the nation's largest water desalination plant. The water from this plant costs $2257 per acre-foot for the first 48000 acre-feet and about $2000 for any additional acre-foot. The county signed a contract with this plant to purchase the first 48000 acre-feet of water for the next 30 years.<ref>Carlsbad Desalination Plant. “Project Overview”. January 2016.</ref> Also, there is another desalination plant in Tampa that produces about 25 million gallons of water per day, and over 2000 other public and industrial plants with production capacity of over 300000 gallons per day. When conducting the economic analysis our plant, such numbers presented by other plants must be considered, so our project can yield positive profit.<ref>Leven, Rachel. “U.S. Desalination Industry Grows Since 2000; See as Essential to Meeting Supply Needs”. Bloomberg BNA. August 21, 2013.</ref>

The profitability of the project is naturally heavily dependent on the price of water in California. If the drought ends, water supply will increase and the water price will drop significantly, but if the drought continues, the price of water in California could rise even more rapidly and allow desalination plants to be more profitable.<ref>Southern California Public Radio. “Where is California water use decreasing?”. December 2015.</ref> It may be difficult to make this plant profitable without a guaranteed purchase or other subsidy from the local governments because the presence of pre-existing plants maintains the water price relatively low. However, external sources of profit such as guaranteed purchase or subsidy are expected because the ramifications of not having desalination plant or other sources of water may be significantly more expensive than the subsidizing a desalination plant assuming the drought continues.<ref>Weiser, Matt. “Could desalination solve California’s water problem?”. The Sacramento Bee. October 18, 2014.</ref>

=Design Basis and Technical Approach=

[[File:Figure_2.jpg|left|200px|frame|Figure 2. Vertical Well Feed]]
Considering the number of pre-existing desalination plant and the potential risk of the drought ending, our plant will be tentatively designed to produce 15 million gallons of drinkable water per day to homes and businesses in the Los Angeles county area. With a recovery rate of near 30%, and a requirement for diluting any brine, the input feed flowrate will be 70.5 million gallons per day. Following the similarly built Carlsbad plant, we expect the plant to last 30 years. The project building timeline is one year based on our relatively small plant size. We will run for 350 days a year while shutting down for cleaning and maintenance once every 3 months (based on DOW standards).<ref name = "Bates">Bates, Wayne. “Cleaning your RO”. Hydranautics. </ref> The plant will be built 500 meters from the coastline on Long Beach. A borehole will be drilled 10 meters into the ground. At this depth there should be seawater aquifers below the coast that regularly recharges itself with seawater from the ocean. A vertical well containing intake holes will be put into the borehole. The intake holes will be equipped with large size filters that can easily be replaced. At the surface, the well will be connected to a large-scale pump in a well pump house. From there, the seawater will be horizontally pumped into a seawater tank in the desalination plant. This will be considered the start of our process. The process ends after pretreatment, desalination, post treatment, and brine treatment are completed. Appendix 1 shows more details on the Design Basis.

The feed composition is assumed to be the standard seawater “average” based on the DOW Chemical ROSA 9.0 Technical Manual in Table 1. However, in reality the seawater composition at the plant site will have to be chemically tested. The final product will be based on L.A. standards for drinking water quality (Table 1). The specifications in the table are upper limits that we cannot exceed. We will aim to reach the salt level standards but any remineralization and pH treatment will be exported and paid for due to the added complexity and cost in making such a system in-house.

Due to the acidity of the desalinated water coming out of the reverse osmosis (RO) system (pH ~ 6), the pumps and pipes carrying the desalinated water for post treatment will likely corrode over time. We have decided to place the desalination plant by the Long Beach Water Reclamation Plant (about 50 meters away) to prevent pipe corrosion.

[[File:Table_1.jpg|center|200px|frame|Table 1. Concentration of the standard seawater, the output from the RO system, and LA standard.]]

=Desalination Overview and Alternatives=

We first began our research into the current landscape of the desalination industry, particularly of plants that were built in the United States. We discovered there were many processes but the main ones included: 
* Membrane Separation (Forward Osmosis and Reverse Osmosis)
* Thermal Separation (Multi Stage Flash Distillation)
* Phase Change Separation (Freeze Separation)
Each process was analyzed for their pros and cons. For our project design, we decided to go with reverse osmosis as our desalination process. This is because it is the most commonly used process in the industry, and for good reason. Unlike thermal separation, which requires huge amounts of energy to supply heat, reverse osmosis is very energy efficient. This reduces energy costs that would be spread across the plant’s lifetime.<ref name = "Republic">Republic of Mauritius NY. "Technical Aspects of Desalination Plant". Sustainable Sanitation and Water Management. </ref>

Forward osmosis is a newer technology being developed that also uses membranes to separate water and salt and may be more energy efficient overall than reverse osmosis. This is because forward osmosis doesn’t require a pump to pressurize seawater through the membranes, but rather leaves a higher salt concentrated solution (AKA draw solution) on the other side of the membrane for the seawater to be drawn into. Then a second step is required to regenerate the concentrated draw solution and produce purified water.<ref>Mallinson, Alissa. "Study Shows Forward Osmosis Desalination Not Energy Efficient." MIT News [Boston]. July 23, 2014.</ref> We decided to move away from forward osmosis because it is a process that is still being refined and developed for desalination and lacks the data, programs, and technologies that reverse osmosis users benefit from. Additionally, a recent study from MIT has claimed that due to the inherent nature of forward osmosis requiring that seawater be drawn to a higher concentrated salt solution before the water and salt is separated, it cannot possibly be more energy efficient than reverse osmosis.<ref>O'Hern, Sean C., Karnik, Rohit. Nanofiltration across Defect-Sealed Nanoporous Monolayer Graphene. American Chemical Society 15.5 (2014); 3254-260. </ref>

Other processes such as freeze separation were too restrictive in their feed and process demands and required a lot of capital. Ultimately, reverse osmosis is simply the best process that exists for desalination and has been ranked the best when performance is evaluated in terms of technical, environmental, and economic aspects. It is a process that is flexible in water quantity, operation start-up and shut-off, and site location.

Nevertheless there are disadvantages to reverse osmosis that must be considered. The feed seawater must not be too salty and too low quality. The pressure to pump the seawater is enormous. The construction time is long for large plants. The capital costs are also high when compared with other processes such as thermal separation. But all of these disadvantages are ultimately outweighed by the advantages reverse osmosis provides.<ref name = "Republic"/>

=Process Overview and Alternatives=

See Appendix 2 (Figure 1) for the Process Flow Diagram (PFD). The process itself is detailed below. The stream flow rates are in Appendix 2 (Table 1). We assumed 10 meters of pipe for every stream except the inlet and outlet streams of the entire PFD.

==Pretreatment==

After the 10 meter vertical well pumps seawater from the Long Beach coast, the feed is pumped 500 meters to a seawater tank that makes it easy to shut-down the system and reduce the large volume of flow. The level control keeps the seawater from overfilling the tank. From there, part of the seawater goes to brine treatment while the rest go to pretreatment. The purpose of pretreatment in the desalination process is to prevent equipment damage by removing particles and microorganisms from the water. A variety of pollutants -- including calcium carbonate, sulfates, silica, clay, bacteria, and other colloids and microorganisms -- can disrupt the reverse osmosis process. Media filtration uses fine grains such as sand or anthracite to filter the particles. We used a large filter to catch larger particle material (e.g. sand). However, simple filtration is not enough to prevent colloid fouling. An ultrafilter (i.e. microfiltration) is used to remove virtually all suspended matter, effectively preventing colloid fouling. The costs associated with ultrafiltration are not prohibitive in terms of the project’s budget.<ref name = "Desalination">Desalination Pre-treatment. Lenntech.</ref>

Reverse osmosis systems can also be disrupted by the growth of microbiological films of bacteria on the film and pipe walls. They are typically treated as colloids during pretreatment, and most organisms are removed from the water by ultrafiltration, but this does not remove the smallest of organic particles. DBNPA is the most effective biocide, but these aren’t often used and have to be removed from the water. Chlorination is an effective way to remove bacteria, but it can cause degradation of the membrane, and must be removed before the reverse osmosis occurs. In our process we will use UV irradiation to remove any organic material that passes through the microfilters.<ref name = "Desalination"/>

Scaling can reduce the flux across the membrane and can build up in pipe corners, causing backup. It is caused by inorganic salts -- specifically sulfates and calcium and magnesium carbonate -- present in the water. Scaling can be prevented and/or controlled with acids, antiscalants, or ion exchange softening. Antiscalants are good for removing sulfates, but not much else. Acidification will only be effective at removing CaCO3 from water but, the decreased pH may deteriorate the membrane and pipes, so it may require future research.<ref name = "Desalination"/> Ion exchange is very expensive and requires more capital to undergo, but is overall the more effective option.

The ion exchanger is a pretreatment system which acts as a method to fight membrane scaling in reverse osmosis. The ion exchanger works by using sodium carboxylate as a resin which will react with Chlorine and Magnesium ions, removing them from the feed. The sodium is substituted out and enters the feed stream, increasing the overall sodium level and TDS. The lack of magnesium and chlorine ions means that there is nothing to react with the carbonates to form scaling salts on the membranes.

==Desalination (Reverse Osmosis)==

Two pumps (P-105 & P-106) take the adjusted feed seawater up to 58 bar. However, before this pumping occurs, about 70% of the original stream gets separated and heads into a PX Pressure Exchanger. This separated stream later joins together with the pumped stream to enter the reverse osmosis system at the required pressure of 58 bar. The streams can combine together as they have equal pressure. It is the concentrate stream from the output of the reverse osmosis system that transfers its energy to the adjusted feed in the PX Pressure Exchanger.

[[File:Figure_3.jpg|left|300px|frame|Figure 3. Inside look of pressure exchanger]]
The pressure of the concentrate stream is about 57 bar and it is fed into an exchanger which works by using vessel that resembles a small pipe within a rotating wheel. The low pressure (LP) feed enters the vessel at 1 bar (see Figure 3). The feed turns and is pushed out into an outlet pipe by the high pressure (HP) concentrate stream at ~57 bar. The feed is now a HP stream via this exchanged energy. The now LP concentrate stream turns and is pushed out into an outlet pipe by the LP feed entering the vessel.

98% of the energy within the concentrate stream can be recovered via the PX Pressure Exchanger.<ref>Desalination Product and Applications Catalog 2013. “Advance Technologies for Seawater Reverse Osmosis and Brackish Water Applications”. Energy Recovery, Inc. 2012-2014.</ref >Thus, the now HP feed stream is not quite 58 bar and will need to be pumped by a small booster pump (P-107).

When designing a reverse osmosis system, we need to make decisions on what specific kind of reverse osmosis system will be implemented. When making the decision, we investigated the advantages of different types of reverse osmosis among the following options: batch vs. continuous process, number of stages, and number of elements.<ref>FILMTEC Reverse Osmosis Membranes. “Water & Process Solutions Technical Manual.” Dow Chemical Company.</ref> Our decision was made based on process optimization which is described later in this report. The end result of that optimization was to go with a single stage, continuous process with five elements in series.

=Post Treatment=

The permeate is sent to post treatment after the salt content is analyzed and controlled if too high or too low due to any unexpected errors. As stated above, the post treatment process will be done 50 meters from our plant by the Long Beach Water Reclamation Plant. The reason we aren’t performing post treatment “in-house” is due to the increased cost it would take for us to build and run the equipment for post treatment. We have the option to use the government’s water treatment facilities that already exist and our 15 million gallons per day output is a relatively small addition to the large Long Beach Water Reclamation Plant. We will assume any costs to post treatment will be pushed to the consumer in the price of purchasing the drinking water from the Long Beach Water Reclamation Plant.

Post treatment at any facility works to bring water to “taste standards”. The water that is the permeate of our process is already meeting L.A. drinking standards as shown above in Table 1 after our mass balance calculations around the reverse osmosis system (Appendix 3, Table 1). The post treatment increases the pH from 6.3 to 7.3 with the addition of base and then adds minerals like fluoride, calcium, and magnesium back in to improve taste.

==Brine Treatment and Byproduct Waste==

A natural consequence of a reverse osmosis procedure is that a lot of salt will be leftover from the membrane separation in a very concentrated form. This is called “brine” and will have to be taken care of. According to EPA requirements, the salt concentration of water entering the environment (via land or sea) cannot be higher than 40,000 ppm.8 The salt concentration of our brine is much higher (around 70,000 ppm). Thus, the brine must be removed or treated. We decided to treat the brine by diluting it as removing brine can be difficult and may require lots of land (e.g. evaporation ponds).

After the concentrate exchanges energy in the energy exchanger, it becomes a low pressure stream. It is pumped into a dilution tank and mixes with 18.2 million gallons per day of seawater from the inlet to become diluted brine which is pumped 600 meters out to sea.

The waste byproducts are the solid buildup on the filters and membranes. The filters and membranes also have a set lifetime and thus will also need to be removed. The waste is all non-hazardous and thus can be put in the municipal dump.

=Process Details and Equipment Sizing=

==Filters and UV Treatment [Pretreatment]==
The cost of our ultrafiltration system is estimated using data from the University of Tennessee at Knoxville.<ref>Qiang He and R. Bruce Robinson Membrane Filtration Systems. “Membrane Module Costs.” University of Tennessee at Knoxville. December 2014.</ref> It estimated the cost of the cost of the ultrafiltration $0.057 per 1000 gallons and the construction and capital costs are $0.39 per gallon per day. The cost of UV treatment is based off of units available from buyultraviolet.com for units that can treat about 450 gallons per minute.<ref>Atlantic Ultraviolet Corporation. “Megatron Water Disinfection System.”</ref> Due to the water having to be in relatively small pipes for treatment we do not expect the costs to scale very much, so we believe these units to be an accurate cost estimate.

==Ion Exchanger [Pretreatment] ==

The ion exchanger was sized using data from DOW. Our process will use 60 vessels with ion exchange beds of volume 5.73 m3. With a flow rate of 30 bed volumes/hour (10,312.2 m3/hr) the system of beds has a capacity of 125% of the feed. This allows for periodic regeneration of the resin, a weak acid cation resin (DOW MAC-3 Resin). This weak cation resin will remove cationic calcium and magnesium from the water to prevent scaling. The mass flow was calculated via a mass balance on the ion exchanger (Appendix 3).

==Reverse Osmosis System [Desalination]==

Each of the 2607 Reverse Osmosis SW30HRLE-440i [Seawater 30 High Rejection Low Energy 440 ft2 interlocking cap membrane] units have the same dimensions. A, the net length, is 1016 mm. B, the total length including the endcap, is 1029 mm. C, the diameter of the filter element, is 201mm. D, the diameter of the permeate pathway, is 29 mm. For a single stage, we have five of the elements in series so the overall length of a single stage is expected to be roughly 5145 mm. The membrane material is polyamide thin film composite. It will last three years.
Operating Conditions [ROSA 9.0]: 
* Temperature: 30oC 
* Pressure: 58 bar 
* Flow Rate: 52.3 Million gal per day 
* Recovery: 28.7%

[[File:Figure_4.png|center|400px|frame|Figure 4. Diagram of Dow FILMTEC SW30HRLE-440i]]
Cleaning Schedule: 
* Conservative Estimate: Once every three months (three days available for cleaning)<ref name = "Bates"/> 
* DOW FILMTEC Cleaning Conditions 
** The normalized permeate flow drops 10%
* The normalized salt passage increases 5 - 10%

==Seawater Holding Tank and Dilution Tank [Brine Treatment]==

The initial seawater holding tank works as a buffer to prevent surges and help perform shutdowns quickly and easily. We expect to run continuously for 350 days a year with the rest of the available time spread out and used for shutting down the plant for cleaning membranes, filters, and the fixing of unexpected errors. There is a big tradeoff between having a holding tank with high retention time and having a holding tank with low cost. With that in mind, we erred on the side of lowering capital cost and went with a retention time of one hour. The volume could be found by multiplying the retention time with the feed flow rate (Stream 1). We assumed that a good height to width ratio for a large water tank would be 4:1.

A mass balance was done around the dilution tank to find what amount of seawater would be necessary to dilute the brine to 45,000 ppm salinity (Appendix 3; Table 3). With this the feed flow rate into the tank could be found. It was found that it takes about 20 minutes of retention time for good dilution of the brine in seawater.8 We again assumed that a good height to width ratio for a large water tank would be 4:1.

We assumed a corrosion allowance of 4 mm wall thickness. We calculated wall thickness using vessel pressure, hoop stress, longitudinal stress, and weld factor. Each tank is made with SS316 to protect against corrosion.<ref>Towler G, Sinnott R. Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013.</ref>

<center>
'''Table 2.''' Tank sizes.
[[File:Table_2.jpg|center|600px]]
</center>

=Process Optimization=
<center> '''Table 3.'''Optimization Criteria
[[File:Table_3.jpg|center|400px]]
</center>
In order to optimize the reverse osmosis system, numerous simulations were run. Three representative cases are reported as shown above (Table 3). The objective of the optimization was to minimize the required water price to break even in 10 years. In order to achieve this goal, the balance between the capital cost from the number of the filter elements and utility cost from energy consumption was weighed. Before the energy recovery system was implemented, the general trend was that if we have more filter elements in series, overall recovery increases but recovery rate per filter decreases, and we need more filter elements. While increasing the elements in series increase the required number of filters, and thereby increases the capital cost, we require less energy to produce the same amount of water. After the energy recovery system is implemented, the correlation between the number of elements in series and the required energy becomes more complicated because high pressure concentrate is used to pressurize the feed. However, the trade-off between the utility cost and capital cost still exists. Of the three cases shown above, single stage, five elements RO system turned out to be the most optimal solution because single stage, single element system required higher utility cost and two stage system slightly did not meet the LA standard. Detailed information which was used for the optimization is available in Appendix 3 and in Appendix 6.

=Ethical, Environmental, and Societal Considerations=

There are a few key considerations we have followed to meet ethical, environmental, and societal guidelines and requirements: 
* Meet L.A. Drinking Water Standards
** Correct ion concentration
* Moved near water treatment plant
** Pipes will not corrode from low pH
* Working below pressure limits
* Water will be locally served and meet taste standards
** Via Post Treatment mineralization
* Follows EPA Environmental Salinity Requirements
* Reduced Energy Requirements
** PX Pressure Exchanger
** Optimized RO System
** Reverse Osmosis is itself Energy Saving

For further environmental considerations, see effluent summary in Appendix 4.

=Economic Analysis=

The pricing of the majority of the process was completed using the Aspen Economic Analyzer, however we had determine costs for the ultrafiltration, reverse osmosis membranes, and the ultraviolet treatment was completed separately. This led to a total ISBL of $56 MM, OSBL of $22.4 MM, Engineering Costs of $22.4 MM and a Contingency of $7.8 MM. We also chose a working Capital of $10 MM. This working capital is sufficient because we have no raw materials costs. We also had FCOP of $5.2 MM and VCOP of $3.5 MM.

We originally were planning on pricing our water at $700 per acre-foot, but after doing economic analysis we had a 10 year NPV of $-76.2 MM. After determining did a sensitivity analysis and determined that if our water was priced at $2104 per acre-foot we could have an NPV of $0 after 10 years and if we were able to get the same price of $2257 per acre-foot as the plant in Carlsbad we would have a 10 year NPV of $8.3 MM which increases to $41.1 MM at 20 years. For more detailed data about the economic analysis see Appendix 6.

=Conclusion and Recommendations=

Our mission was to create an economically viable process that will desalinate water for human consumption in the L.A. County. We have found after research and optimization that the best process to undergo desalination is a continuous single stage, five elements reverse osmosis system. Our process can meet the required L.A. drinking water standard and we have set up plans for post treatment at a local water reclamation site. While meeting the drinking water standards, we also met our ethical, environmental, and social standards.

Nevertheless, at a sale price of $700 per acre-foot, our proposed desalination plant in Los Angeles, California looks to be a very expensive project to undergo. Without any sort of government assistance to carry the heavy capital costs of such a project, we would suggest the plant not be built. However, after performing a sensitivity analysis at higher (and less conservative) price points we reached a breakeven point in 10 years with a sale price of $2104 per acre-foot of drinking water. Research has shown that the droughts striking California look only to get worse in the near future. With limited freshwater resources, there is no option other than to consider desalination. Therefore it is reasonable to expect we could get near the same price deal as local desalination plants like the Carlsbad Plant ($2257 per acre-foot) As we have met all design constraints for this project, we hope to convince policymakers to adopt our project at least $2104 per acre-foot water. If this deal is reached, we recommend that this project and design be funded.

However, there are ways to improve our design with future work. Right now all the pipes diameters are based on having a flowrate of 1.5 m/s. But in reality, we’d need to use standard piping sizes to get more accurate pricing and energy costs. Also, in reality we would be purchasing some of these vessels in bulk quantities (i.e. bulk filters) so the price could be lowered with negotiation and economies of scale. We also have pumps that are providing very little pressure for their high cost. We can remove these pumps and simply run at higher pressures earlier. We can potentially reduce early energy costs by building a “floating” desalination plant in the sea to reduce pumping distance from the well.

=Appendices=

==Appendix 1: Design Basis==
[[File:Appendix_1.png|thumb|border|center|800px|middle|]]

==Appendix 2: Process Flow Diagram==
[[File:Appendix_2.png|center]]
==Appendix 3: Mass and Energy Balance Tables==
'''Table 1'''. Mass balance on 1 stage 5 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T1.png|center|500px|]]

'''Table 2'''. Mass balance on 1 stage 1 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T2.png|center|500px|]]

'''Table 3'''. Mass balance on 2 stage 5,5 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T3.png|center|600px|]]

'''Table 4'''. 1 stage 5 elements CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T4.png|center|300px|]]

'''Table 5'''. 1 stage 1 element CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T5.png|center|300px|]]

'''Table 6'''. 2 stages 5,5 elements CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T6.png|center|300px|]]

'''Table 7'''. Energy balance on 1 stage 5 elements RO system
[[File:Appendix_3_T7.png|center|800px|]]

==Appendix 4: Effluent Summary==
[[File:Appendix_4.png|center]]

==Appendix 5: Cost of Production==
[[File:Appendix_5.png|center]]

==Appendix 6: Economic Anlaysis==
===Case I ===
Water is $700; RO with 1 stage and 5 elements
[[File:Appendix_6_C1.png|center]]

===Case II===
Water is $1400; RO with 1 stage and 5 elements
[[File:Appendix_6_C2.png|center]]

===Case III: Carlsbad ===
Water is $2257; RO with 1 stage and 5 elements
[[File:Appendix_6_C3.png|center]]

===Case IV: Breakeven (NPV = 0 at 10 years)===
Water is $2104; RO with 1 stage and 5 elements
[[File:Appendix_6_C4.png|center]]

===Case V: Breakeven (NPV = 0 at 10 years)===
Water is $2112; RO with 2 stages and 5,5 elements
[[File:Appendix_6_C5.png|center]]

===Case VI: Breakeven (NPV = 0 at 10 years)===
Water is $2138 with 1 stage and 1 element
[[File:Appendix_6_C6.png|center]]

==Appendix 7: ROSA Simulation Outputs==

'''Table 1'''. System detail of 1 stage 5 elements RO system
[[File:Appendix_7_T1.png|center]]

'''Table 2'''. System detail of 1 stage 1 elements RO system
[[File:Appendix_7_T2.png|center]]

'''Table 3'''. System detail of 2 stage 5,5 elements system
[[File:Appendix_7_T3.png|center]]

'''Table 4'''. 1 stage 5 elements RO system concentrations and pH in each stream
[[File:Appendix_7_T4.png|center]]

'''Table 5'''. 1 stage 1 element RO system concentration and pH in each stream
[[File:Appendix_7_T5.png|center]]

'''Table 6'''. 2 stage 5,5 elements RO system concentrations and pH in each stream
[[File:Appendix_7_T6.png|center]]

=References=
<references/>

Desalination - Team E

2016-03-12T03:34:42Z

Bts511: /* Appendix 3: Mass and Energy Balance Tables */

Group E Corporation

Authors: Hassan Ali, Woo Soo Choe, Brett Sleyster, Jake Stolley

Instructors: Fengqi You, David Wegerer

March 11, 2016

=Executive Summary=

Outlined in this report is the project model report for the proposed desalination plant in Los Angeles, California which will provide drinking water to local citizens using reverse osmosis process. This location was chosen due to the region’s propensity for droughts in recent years. The fresh water resources in California have plummeted as water prices have risen 3-4% each year. This is only expected to continue. Other companies have worked deals with the Californian government to create desalination plants that provide a new fresh water drinking source. This suggests that there is a strong market here with potential growth.

The plant will have an output of 15 million gallons per day, which is about medium size relative to other plants. Thus the build time will be about one year. The plant will run for 30 years, with 350 days per year of use. Shutdown time will be used for cleaning and maintenance of equipment. The plant will be built on Long Beach, 500 meters from the coast nearby the Long Beach Water Reclamation Plant (LBWP). The input will withdraw 70.5 million gallons per day of seawater that will be used for reverse osmosis and dilution of brine.

Reverse osmosis was chosen as the desalination method for its low energy cost and flexibility. We designed a process that will pretreat the seawater to remove waste, colloids, bacteria, and membrane-damaging scalants. Once this seawater is treated is pumped to heavy pressures through a reverse osmosis system. After doing process optimization under various system parameter changes, we found a continuous, single stage, five element reverse osmosis process to be best. It produces a near 30% recovery. We reduce energy costs by including a pressure exchanger between our concentrate outlet and adjusted feed into the reverse osmosis system.

After performing mass and energy balances, we were able to size equipment and find our permeate composition. It turns out our process will meet L.A. standards for drinking water quality. However, post treatment will be done by the local LBWP to make the water taste better and be non-caustic to pipes. This will reduce “in-house” costs. Many environmental considerations were made, to meet our environmental standards. We meet EPA salinity requirements for our brine waste by diluting it. We also don’t use a lot of energy and thus reduce our fossil fuel consumption and CO2 generation. To be safe, we work well under max pressure limits for our equipment. We make sure we are ethical and meet all standards that our local county requires of us.

After an economic evaluation at a conservative well-water price of $700 per acre-foot, we realize that our revenue is way too low to cover our costs. Capital costs are especially large with regards to our large water tanks. We find that we reach a 10 year NPV breakpoint of $2104 per acre-foot after performing a sensitivity analysis. The Carlsbad Plant in San Diego worked out a deal for $2257 per acre-foot.

As we have met all design constraints for this project, we hope to convince policymakers to adopt our project at least $2104 per acre-foot water. If this deal is reached, we recommend that this project and design be funded. There is future work to be done to improve costs such as performing pump optimization. However, this work is minimal and won’t harm our bottom line if we reach our proposed price points.

__TOC__

=Introduction=

Due to a combination of various factors--including rising human population, climate change, increased energy consumption, and land erosion--the availability of fresh water as a resource for human consumption has been dwindling. This has had a negative impact on the world, but this impact has been especially felt in areas with natural shortages in fresh water. Desert regions, such as Saudi Arabia, have resorted to desalinating seawater at the coast and transporting up to their cities. This encompasses their entire freshwater system as very few natural water sources remain. In the United States there is a region which has shown increasing signs of following in the footsteps of Saudi Arabia and becoming fully reliant on desalination: Los Angeles, California.

=Location and Market Analysis=

In this project, Los Angeles was chosen as the location our plant, because the area has been subject to a severe drought for the last four years, as shown in Figure 1. The rapid decrease in the available water has caused permanent damages to the ground, decreased the amount of water in the reservoir from 45% to 25%, and reduced the amount of hydroelectric power produced in the state. Such deficiency in usable water has increased the water prices in Los Angeles by 6-7% annually for the last decade and is expected to rise by 3.4% annually for typical customers for the next five years. For customers who use more than 20000 gallons a year, the bill is expected to increase by 34% by 2021. This price increase is occurring because of both drought and maintenance. <ref>Pacific Institute: The California Drought. 2014.</ref>

[[File:Figure_1.jpg|right|200px|frame|Figure 1. California Drought Index.]]
As of March 2015 the Metropolitan Water District of Southern California was willing to pay as much as $700 per acre-foot of water to farmers in the Sacramento Valley, so we assume this will be the target price of the water produced by the this desalination plant.<ref>Hessel, Phil. “Dry Southern California Offers Northern Farmers Top Dollar for Water”. NBC News. March 18, 2015. </ref>

Currently, there is a plant in Carlsbad which supplies drinking water to the San Diego area. The plant can produce up to 50 million gallons per day and is the nation's largest water desalination plant. The water from this plant costs $2257 per acre-foot for the first 48000 acre-feet and about $2000 for any additional acre-foot. The county signed a contract with this plant to purchase the first 48000 acre-feet of water for the next 30 years.<ref>Carlsbad Desalination Plant. “Project Overview”. January 2016.</ref> Also, there is another desalination plant in Tampa that produces about 25 million gallons of water per day, and over 2000 other public and industrial plants with production capacity of over 300000 gallons per day. When conducting the economic analysis our plant, such numbers presented by other plants must be considered, so our project can yield positive profit.<ref>Leven, Rachel. “U.S. Desalination Industry Grows Since 2000; See as Essential to Meeting Supply Needs”. Bloomberg BNA. August 21, 2013.</ref>

The profitability of the project is naturally heavily dependent on the price of water in California. If the drought ends, water supply will increase and the water price will drop significantly, but if the drought continues, the price of water in California could rise even more rapidly and allow desalination plants to be more profitable.<ref>Southern California Public Radio. “Where is California water use decreasing?”. December 2015.</ref> It may be difficult to make this plant profitable without a guaranteed purchase or other subsidy from the local governments because the presence of pre-existing plants maintains the water price relatively low. However, external sources of profit such as guaranteed purchase or subsidy are expected because the ramifications of not having desalination plant or other sources of water may be significantly more expensive than the subsidizing a desalination plant assuming the drought continues.<ref>Weiser, Matt. “Could desalination solve California’s water problem?”. The Sacramento Bee. October 18, 2014.</ref>

=Design Basis and Technical Approach=

[[File:Figure_2.jpg|left|200px|frame|Figure 2. Vertical Well Feed]]
Considering the number of pre-existing desalination plant and the potential risk of the drought ending, our plant will be tentatively designed to produce 15 million gallons of drinkable water per day to homes and businesses in the Los Angeles county area. With a recovery rate of near 30%, and a requirement for diluting any brine, the input feed flowrate will be 70.5 million gallons per day. Following the similarly built Carlsbad plant, we expect the plant to last 30 years. The project building timeline is one year based on our relatively small plant size. We will run for 350 days a year while shutting down for cleaning and maintenance once every 3 months (based on DOW standards).<ref name = "Bates">Bates, Wayne. “Cleaning your RO”. Hydranautics. </ref> The plant will be built 500 meters from the coastline on Long Beach. A borehole will be drilled 10 meters into the ground. At this depth there should be seawater aquifers below the coast that regularly recharges itself with seawater from the ocean. A vertical well containing intake holes will be put into the borehole. The intake holes will be equipped with large size filters that can easily be replaced. At the surface, the well will be connected to a large-scale pump in a well pump house. From there, the seawater will be horizontally pumped into a seawater tank in the desalination plant. This will be considered the start of our process. The process ends after pretreatment, desalination, post treatment, and brine treatment are completed. Appendix 1 shows more details on the Design Basis.

The feed composition is assumed to be the standard seawater “average” based on the DOW Chemical ROSA 9.0 Technical Manual in Table 1. However, in reality the seawater composition at the plant site will have to be chemically tested. The final product will be based on L.A. standards for drinking water quality (Table 1). The specifications in the table are upper limits that we cannot exceed. We will aim to reach the salt level standards but any remineralization and pH treatment will be exported and paid for due to the added complexity and cost in making such a system in-house.

Due to the acidity of the desalinated water coming out of the reverse osmosis (RO) system (pH ~ 6), the pumps and pipes carrying the desalinated water for post treatment will likely corrode over time. We have decided to place the desalination plant by the Long Beach Water Reclamation Plant (about 50 meters away) to prevent pipe corrosion.

[[File:Table_1.jpg|center|200px|frame|Table 1. Concentration of the standard seawater, the output from the RO system, and LA standard.]]

=Desalination Overview and Alternatives=

We first began our research into the current landscape of the desalination industry, particularly of plants that were built in the United States. We discovered there were many processes but the main ones included: 
* Membrane Separation (Forward Osmosis and Reverse Osmosis)
* Thermal Separation (Multi Stage Flash Distillation)
* Phase Change Separation (Freeze Separation)
Each process was analyzed for their pros and cons. For our project design, we decided to go with reverse osmosis as our desalination process. This is because it is the most commonly used process in the industry, and for good reason. Unlike thermal separation, which requires huge amounts of energy to supply heat, reverse osmosis is very energy efficient. This reduces energy costs that would be spread across the plant’s lifetime.<ref name = "Republic">Republic of Mauritius NY. "Technical Aspects of Desalination Plant". Sustainable Sanitation and Water Management. </ref>

Forward osmosis is a newer technology being developed that also uses membranes to separate water and salt and may be more energy efficient overall than reverse osmosis. This is because forward osmosis doesn’t require a pump to pressurize seawater through the membranes, but rather leaves a higher salt concentrated solution (AKA draw solution) on the other side of the membrane for the seawater to be drawn into. Then a second step is required to regenerate the concentrated draw solution and produce purified water.<ref>Mallinson, Alissa. "Study Shows Forward Osmosis Desalination Not Energy Efficient." MIT News [Boston]. July 23, 2014.</ref> We decided to move away from forward osmosis because it is a process that is still being refined and developed for desalination and lacks the data, programs, and technologies that reverse osmosis users benefit from. Additionally, a recent study from MIT has claimed that due to the inherent nature of forward osmosis requiring that seawater be drawn to a higher concentrated salt solution before the water and salt is separated, it cannot possibly be more energy efficient than reverse osmosis.<ref>O'Hern, Sean C., Karnik, Rohit. Nanofiltration across Defect-Sealed Nanoporous Monolayer Graphene. American Chemical Society 15.5 (2014); 3254-260. </ref>

Other processes such as freeze separation were too restrictive in their feed and process demands and required a lot of capital. Ultimately, reverse osmosis is simply the best process that exists for desalination and has been ranked the best when performance is evaluated in terms of technical, environmental, and economic aspects. It is a process that is flexible in water quantity, operation start-up and shut-off, and site location.

Nevertheless there are disadvantages to reverse osmosis that must be considered. The feed seawater must not be too salty and too low quality. The pressure to pump the seawater is enormous. The construction time is long for large plants. The capital costs are also high when compared with other processes such as thermal separation. But all of these disadvantages are ultimately outweighed by the advantages reverse osmosis provides.<ref name = "Republic"/>

=Process Overview and Alternatives=

See Appendix 2 (Figure 1) for the Process Flow Diagram (PFD). The process itself is detailed below. The stream flow rates are in Appendix 2 (Table 1). We assumed 10 meters of pipe for every stream except the inlet and outlet streams of the entire PFD.

==Pretreatment==

After the 10 meter vertical well pumps seawater from the Long Beach coast, the feed is pumped 500 meters to a seawater tank that makes it easy to shut-down the system and reduce the large volume of flow. The level control keeps the seawater from overfilling the tank. From there, part of the seawater goes to brine treatment while the rest go to pretreatment. The purpose of pretreatment in the desalination process is to prevent equipment damage by removing particles and microorganisms from the water. A variety of pollutants -- including calcium carbonate, sulfates, silica, clay, bacteria, and other colloids and microorganisms -- can disrupt the reverse osmosis process. Media filtration uses fine grains such as sand or anthracite to filter the particles. We used a large filter to catch larger particle material (e.g. sand). However, simple filtration is not enough to prevent colloid fouling. An ultrafilter (i.e. microfiltration) is used to remove virtually all suspended matter, effectively preventing colloid fouling. The costs associated with ultrafiltration are not prohibitive in terms of the project’s budget.<ref name = "Desalination">Desalination Pre-treatment. Lenntech.</ref>

Reverse osmosis systems can also be disrupted by the growth of microbiological films of bacteria on the film and pipe walls. They are typically treated as colloids during pretreatment, and most organisms are removed from the water by ultrafiltration, but this does not remove the smallest of organic particles. DBNPA is the most effective biocide, but these aren’t often used and have to be removed from the water. Chlorination is an effective way to remove bacteria, but it can cause degradation of the membrane, and must be removed before the reverse osmosis occurs. In our process we will use UV irradiation to remove any organic material that passes through the microfilters.<ref name = "Desalination"/>

Scaling can reduce the flux across the membrane and can build up in pipe corners, causing backup. It is caused by inorganic salts -- specifically sulfates and calcium and magnesium carbonate -- present in the water. Scaling can be prevented and/or controlled with acids, antiscalants, or ion exchange softening. Antiscalants are good for removing sulfates, but not much else. Acidification will only be effective at removing CaCO3 from water but, the decreased pH may deteriorate the membrane and pipes, so it may require future research.<ref name = "Desalination"/> Ion exchange is very expensive and requires more capital to undergo, but is overall the more effective option.

The ion exchanger is a pretreatment system which acts as a method to fight membrane scaling in reverse osmosis. The ion exchanger works by using sodium carboxylate as a resin which will react with Chlorine and Magnesium ions, removing them from the feed. The sodium is substituted out and enters the feed stream, increasing the overall sodium level and TDS. The lack of magnesium and chlorine ions means that there is nothing to react with the carbonates to form scaling salts on the membranes.

==Desalination (Reverse Osmosis)==

Two pumps (P-105 & P-106) take the adjusted feed seawater up to 58 bar. However, before this pumping occurs, about 70% of the original stream gets separated and heads into a PX Pressure Exchanger. This separated stream later joins together with the pumped stream to enter the reverse osmosis system at the required pressure of 58 bar. The streams can combine together as they have equal pressure. It is the concentrate stream from the output of the reverse osmosis system that transfers its energy to the adjusted feed in the PX Pressure Exchanger.

[[File:Figure_3.jpg|left|300px|frame|Figure 3. Inside look of pressure exchanger]]
The pressure of the concentrate stream is about 57 bar and it is fed into an exchanger which works by using vessel that resembles a small pipe within a rotating wheel. The low pressure (LP) feed enters the vessel at 1 bar (see Figure 3). The feed turns and is pushed out into an outlet pipe by the high pressure (HP) concentrate stream at ~57 bar. The feed is now a HP stream via this exchanged energy. The now LP concentrate stream turns and is pushed out into an outlet pipe by the LP feed entering the vessel.

98% of the energy within the concentrate stream can be recovered via the PX Pressure Exchanger.<ref>Desalination Product and Applications Catalog 2013. “Advance Technologies for Seawater Reverse Osmosis and Brackish Water Applications”. Energy Recovery, Inc. 2012-2014.</ref >Thus, the now HP feed stream is not quite 58 bar and will need to be pumped by a small booster pump (P-107).

When designing a reverse osmosis system, we need to make decisions on what specific kind of reverse osmosis system will be implemented. When making the decision, we investigated the advantages of different types of reverse osmosis among the following options: batch vs. continuous process, number of stages, and number of elements.<ref>FILMTEC Reverse Osmosis Membranes. “Water & Process Solutions Technical Manual.” Dow Chemical Company.</ref> Our decision was made based on process optimization which is described later in this report. The end result of that optimization was to go with a single stage, continuous process with five elements in series.

=Post Treatment=

The permeate is sent to post treatment after the salt content is analyzed and controlled if too high or too low due to any unexpected errors. As stated above, the post treatment process will be done 50 meters from our plant by the Long Beach Water Reclamation Plant. The reason we aren’t performing post treatment “in-house” is due to the increased cost it would take for us to build and run the equipment for post treatment. We have the option to use the government’s water treatment facilities that already exist and our 15 million gallons per day output is a relatively small addition to the large Long Beach Water Reclamation Plant. We will assume any costs to post treatment will be pushed to the consumer in the price of purchasing the drinking water from the Long Beach Water Reclamation Plant.

Post treatment at any facility works to bring water to “taste standards”. The water that is the permeate of our process is already meeting L.A. drinking standards as shown above in Table 1 after our mass balance calculations around the reverse osmosis system (Appendix 3, Table 1). The post treatment increases the pH from 6.3 to 7.3 with the addition of base and then adds minerals like fluoride, calcium, and magnesium back in to improve taste.

==Brine Treatment and Byproduct Waste==

A natural consequence of a reverse osmosis procedure is that a lot of salt will be leftover from the membrane separation in a very concentrated form. This is called “brine” and will have to be taken care of. According to EPA requirements, the salt concentration of water entering the environment (via land or sea) cannot be higher than 40,000 ppm.8 The salt concentration of our brine is much higher (around 70,000 ppm). Thus, the brine must be removed or treated. We decided to treat the brine by diluting it as removing brine can be difficult and may require lots of land (e.g. evaporation ponds).

After the concentrate exchanges energy in the energy exchanger, it becomes a low pressure stream. It is pumped into a dilution tank and mixes with 18.2 million gallons per day of seawater from the inlet to become diluted brine which is pumped 600 meters out to sea.

The waste byproducts are the solid buildup on the filters and membranes. The filters and membranes also have a set lifetime and thus will also need to be removed. The waste is all non-hazardous and thus can be put in the municipal dump.

=Process Details and Equipment Sizing=

==Filters and UV Treatment [Pretreatment]==
The cost of our ultrafiltration system is estimated using data from the University of Tennessee at Knoxville.<ref>Qiang He and R. Bruce Robinson Membrane Filtration Systems. “Membrane Module Costs.” University of Tennessee at Knoxville. December 2014.</ref> It estimated the cost of the cost of the ultrafiltration $0.057 per 1000 gallons and the construction and capital costs are $0.39 per gallon per day. The cost of UV treatment is based off of units available from buyultraviolet.com for units that can treat about 450 gallons per minute.<ref>Atlantic Ultraviolet Corporation. “Megatron Water Disinfection System.”</ref> Due to the water having to be in relatively small pipes for treatment we do not expect the costs to scale very much, so we believe these units to be an accurate cost estimate.

==Ion Exchanger [Pretreatment] ==

The ion exchanger was sized using data from DOW. Our process will use 60 vessels with ion exchange beds of volume 5.73 m3. With a flow rate of 30 bed volumes/hour (10,312.2 m3/hr) the system of beds has a capacity of 125% of the feed. This allows for periodic regeneration of the resin, a weak acid cation resin (DOW MAC-3 Resin). This weak cation resin will remove cationic calcium and magnesium from the water to prevent scaling. The mass flow was calculated via a mass balance on the ion exchanger (Appendix 3).

==Reverse Osmosis System [Desalination]==

Each of the 2607 Reverse Osmosis SW30HRLE-440i [Seawater 30 High Rejection Low Energy 440 ft2 interlocking cap membrane] units have the same dimensions. A, the net length, is 1016 mm. B, the total length including the endcap, is 1029 mm. C, the diameter of the filter element, is 201mm. D, the diameter of the permeate pathway, is 29 mm. For a single stage, we have five of the elements in series so the overall length of a single stage is expected to be roughly 5145 mm. The membrane material is polyamide thin film composite. It will last three years.
Operating Conditions [ROSA 9.0]: 
* Temperature: 30oC 
* Pressure: 58 bar 
* Flow Rate: 52.3 Million gal per day 
* Recovery: 28.7%

[[File:Figure_4.png|center|400px|frame|Figure 4. Diagram of Dow FILMTEC SW30HRLE-440i]]
Cleaning Schedule: 
* Conservative Estimate: Once every three months (three days available for cleaning)<ref name = "Bates"/> 
* DOW FILMTEC Cleaning Conditions 
** The normalized permeate flow drops 10%
* The normalized salt passage increases 5 - 10%

==Seawater Holding Tank and Dilution Tank [Brine Treatment]==

The initial seawater holding tank works as a buffer to prevent surges and help perform shutdowns quickly and easily. We expect to run continuously for 350 days a year with the rest of the available time spread out and used for shutting down the plant for cleaning membranes, filters, and the fixing of unexpected errors. There is a big tradeoff between having a holding tank with high retention time and having a holding tank with low cost. With that in mind, we erred on the side of lowering capital cost and went with a retention time of one hour. The volume could be found by multiplying the retention time with the feed flow rate (Stream 1). We assumed that a good height to width ratio for a large water tank would be 4:1.

A mass balance was done around the dilution tank to find what amount of seawater would be necessary to dilute the brine to 45,000 ppm salinity (Appendix 3; Table 3). With this the feed flow rate into the tank could be found. It was found that it takes about 20 minutes of retention time for good dilution of the brine in seawater.8 We again assumed that a good height to width ratio for a large water tank would be 4:1.

We assumed a corrosion allowance of 4 mm wall thickness. We calculated wall thickness using vessel pressure, hoop stress, longitudinal stress, and weld factor. Each tank is made with SS316 to protect against corrosion.<ref>Towler G, Sinnott R. Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013.</ref>

<center>
'''Table 2.''' Tank sizes.
[[File:Table_2.jpg|center|600px]]
</center>

=Process Optimization=
<center> '''Table 3.'''Optimization Criteria
[[File:Table_3.jpg|center|400px]]
</center>
In order to optimize the reverse osmosis system, numerous simulations were run. Three representative cases are reported as shown above (Table 3). The objective of the optimization was to minimize the required water price to break even in 10 years. In order to achieve this goal, the balance between the capital cost from the number of the filter elements and utility cost from energy consumption was weighed. Before the energy recovery system was implemented, the general trend was that if we have more filter elements in series, overall recovery increases but recovery rate per filter decreases, and we need more filter elements. While increasing the elements in series increase the required number of filters, and thereby increases the capital cost, we require less energy to produce the same amount of water. After the energy recovery system is implemented, the correlation between the number of elements in series and the required energy becomes more complicated because high pressure concentrate is used to pressurize the feed. However, the trade-off between the utility cost and capital cost still exists. Of the three cases shown above, single stage, five elements RO system turned out to be the most optimal solution because single stage, single element system required higher utility cost and two stage system slightly did not meet the LA standard. Detailed information which was used for the optimization is available in Appendix 3 and in Appendix 6.

=Ethical, Environmental, and Societal Considerations=

There are a few key considerations we have followed to meet ethical, environmental, and societal guidelines and requirements: 
* Meet L.A. Drinking Water Standards
** Correct ion concentration
* Moved near water treatment plant
** Pipes will not corrode from low pH
* Working below pressure limits
* Water will be locally served and meet taste standards
** Via Post Treatment mineralization
* Follows EPA Environmental Salinity Requirements
* Reduced Energy Requirements
** PX Pressure Exchanger
** Optimized RO System
** Reverse Osmosis is itself Energy Saving

For further environmental considerations, see effluent summary in Appendix 4.

=Economic Analysis=

The pricing of the majority of the process was completed using the Aspen Economic Analyzer, however we had determine costs for the ultrafiltration, reverse osmosis membranes, and the ultraviolet treatment was completed separately. This led to a total ISBL of $56 MM, OSBL of $22.4 MM, Engineering Costs of $22.4 MM and a Contingency of $7.8 MM. We also chose a working Capital of $10 MM. This working capital is sufficient because we have no raw materials costs. We also had FCOP of $5.2 MM and VCOP of $3.5 MM.

We originally were planning on pricing our water at $700 per acre-foot, but after doing economic analysis we had a 10 year NPV of $-76.2 MM. After determining did a sensitivity analysis and determined that if our water was priced at $2104 per acre-foot we could have an NPV of $0 after 10 years and if we were able to get the same price of $2257 per acre-foot as the plant in Carlsbad we would have a 10 year NPV of $8.3 MM which increases to $41.1 MM at 20 years. For more detailed data about the economic analysis see Appendix 6.

=Conclusion and Recommendations=

Our mission was to create an economically viable process that will desalinate water for human consumption in the L.A. County. We have found after research and optimization that the best process to undergo desalination is a continuous single stage, five elements reverse osmosis system. Our process can meet the required L.A. drinking water standard and we have set up plans for post treatment at a local water reclamation site. While meeting the drinking water standards, we also met our ethical, environmental, and social standards.

Nevertheless, at a sale price of $700 per acre-foot, our proposed desalination plant in Los Angeles, California looks to be a very expensive project to undergo. Without any sort of government assistance to carry the heavy capital costs of such a project, we would suggest the plant not be built. However, after performing a sensitivity analysis at higher (and less conservative) price points we reached a breakeven point in 10 years with a sale price of $2104 per acre-foot of drinking water. Research has shown that the droughts striking California look only to get worse in the near future. With limited freshwater resources, there is no option other than to consider desalination. Therefore it is reasonable to expect we could get near the same price deal as local desalination plants like the Carlsbad Plant ($2257 per acre-foot) As we have met all design constraints for this project, we hope to convince policymakers to adopt our project at least $2104 per acre-foot water. If this deal is reached, we recommend that this project and design be funded.

However, there are ways to improve our design with future work. Right now all the pipes diameters are based on having a flowrate of 1.5 m/s. But in reality, we’d need to use standard piping sizes to get more accurate pricing and energy costs. Also, in reality we would be purchasing some of these vessels in bulk quantities (i.e. bulk filters) so the price could be lowered with negotiation and economies of scale. We also have pumps that are providing very little pressure for their high cost. We can remove these pumps and simply run at higher pressures earlier. We can potentially reduce early energy costs by building a “floating” desalination plant in the sea to reduce pumping distance from the well.

=Appendices=

==Appendix 1: Design Basis==
[[File:Appendix_1.png|thumb|border|center|800px|middle|]]

==Appendix 2: Process Flow Diagram==
[[File:Appendix_2.png|center]]
==Appendix 3: Mass and Energy Balance Tables==
'''Table 1'''. Mass balance on 1 stage 5 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T1.png|center|500px|]]

'''Table 2'''. Mass balance on 1 stage 1 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T2.png|center|500px|]]

'''Table 3'''. Mass balance on 2 stage 5,5 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T3.png|center|600px|]]

'''Table 4'''. 1 stage 5 elements CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T4.png|center|200px|]]

'''Table 5'''. 1 stage 1 element CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T5.png|center|200px|]]

'''Table 6'''. 2 stages 5,5 elements CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T6.png|center|200px|]]

'''Table 7'''. Energy balance on 1 stage 5 elements RO system
[[File:Appendix_3_T7.png|center|800px|]]

==Appendix 4: Effluent Summary==
[[File:Appendix_4.png|center]]

==Appendix 5: Cost of Production==
[[File:Appendix_5.png|center]]

==Appendix 6: Economic Anlaysis==
===Case I ===
Water is $700; RO with 1 stage and 5 elements
[[File:Appendix_6_C1.png|center]]

===Case II===
Water is $1400; RO with 1 stage and 5 elements
[[File:Appendix_6_C2.png|center]]

===Case III: Carlsbad ===
Water is $2257; RO with 1 stage and 5 elements
[[File:Appendix_6_C3.png|center]]

===Case IV: Breakeven (NPV = 0 at 10 years)===
Water is $2104; RO with 1 stage and 5 elements
[[File:Appendix_6_C4.png|center]]

===Case V: Breakeven (NPV = 0 at 10 years)===
Water is $2112; RO with 2 stages and 5,5 elements
[[File:Appendix_6_C5.png|center]]

===Case VI: Breakeven (NPV = 0 at 10 years)===
Water is $2138 with 1 stage and 1 element
[[File:Appendix_6_C6.png|center]]

==Appendix 7: ROSA Simulation Outputs==

'''Table 1'''. System detail of 1 stage 5 elements RO system
[[File:Appendix_7_T1.png|center]]

'''Table 2'''. System detail of 1 stage 1 elements RO system
[[File:Appendix_7_T2.png|center]]

'''Table 3'''. System detail of 2 stage 5,5 elements system
[[File:Appendix_7_T3.png|center]]

'''Table 4'''. 1 stage 5 elements RO system concentrations and pH in each stream
[[File:Appendix_7_T4.png|center]]

'''Table 5'''. 1 stage 1 element RO system concentration and pH in each stream
[[File:Appendix_7_T5.png|center]]

'''Table 6'''. 2 stage 5,5 elements RO system concentrations and pH in each stream
[[File:Appendix_7_T6.png|center]]

=References=
<references/>

Desalination - Team E

2016-03-12T03:33:57Z

Bts511: /* Appendix 3: Mass and Energy Balance Tables */

Group E Corporation

Authors: Hassan Ali, Woo Soo Choe, Brett Sleyster, Jake Stolley

Instructors: Fengqi You, David Wegerer

March 11, 2016

=Executive Summary=

Outlined in this report is the project model report for the proposed desalination plant in Los Angeles, California which will provide drinking water to local citizens using reverse osmosis process. This location was chosen due to the region’s propensity for droughts in recent years. The fresh water resources in California have plummeted as water prices have risen 3-4% each year. This is only expected to continue. Other companies have worked deals with the Californian government to create desalination plants that provide a new fresh water drinking source. This suggests that there is a strong market here with potential growth.

The plant will have an output of 15 million gallons per day, which is about medium size relative to other plants. Thus the build time will be about one year. The plant will run for 30 years, with 350 days per year of use. Shutdown time will be used for cleaning and maintenance of equipment. The plant will be built on Long Beach, 500 meters from the coast nearby the Long Beach Water Reclamation Plant (LBWP). The input will withdraw 70.5 million gallons per day of seawater that will be used for reverse osmosis and dilution of brine.

Reverse osmosis was chosen as the desalination method for its low energy cost and flexibility. We designed a process that will pretreat the seawater to remove waste, colloids, bacteria, and membrane-damaging scalants. Once this seawater is treated is pumped to heavy pressures through a reverse osmosis system. After doing process optimization under various system parameter changes, we found a continuous, single stage, five element reverse osmosis process to be best. It produces a near 30% recovery. We reduce energy costs by including a pressure exchanger between our concentrate outlet and adjusted feed into the reverse osmosis system.

After performing mass and energy balances, we were able to size equipment and find our permeate composition. It turns out our process will meet L.A. standards for drinking water quality. However, post treatment will be done by the local LBWP to make the water taste better and be non-caustic to pipes. This will reduce “in-house” costs. Many environmental considerations were made, to meet our environmental standards. We meet EPA salinity requirements for our brine waste by diluting it. We also don’t use a lot of energy and thus reduce our fossil fuel consumption and CO2 generation. To be safe, we work well under max pressure limits for our equipment. We make sure we are ethical and meet all standards that our local county requires of us.

After an economic evaluation at a conservative well-water price of $700 per acre-foot, we realize that our revenue is way too low to cover our costs. Capital costs are especially large with regards to our large water tanks. We find that we reach a 10 year NPV breakpoint of $2104 per acre-foot after performing a sensitivity analysis. The Carlsbad Plant in San Diego worked out a deal for $2257 per acre-foot.

As we have met all design constraints for this project, we hope to convince policymakers to adopt our project at least $2104 per acre-foot water. If this deal is reached, we recommend that this project and design be funded. There is future work to be done to improve costs such as performing pump optimization. However, this work is minimal and won’t harm our bottom line if we reach our proposed price points.

__TOC__

=Introduction=

Due to a combination of various factors--including rising human population, climate change, increased energy consumption, and land erosion--the availability of fresh water as a resource for human consumption has been dwindling. This has had a negative impact on the world, but this impact has been especially felt in areas with natural shortages in fresh water. Desert regions, such as Saudi Arabia, have resorted to desalinating seawater at the coast and transporting up to their cities. This encompasses their entire freshwater system as very few natural water sources remain. In the United States there is a region which has shown increasing signs of following in the footsteps of Saudi Arabia and becoming fully reliant on desalination: Los Angeles, California.

=Location and Market Analysis=

In this project, Los Angeles was chosen as the location our plant, because the area has been subject to a severe drought for the last four years, as shown in Figure 1. The rapid decrease in the available water has caused permanent damages to the ground, decreased the amount of water in the reservoir from 45% to 25%, and reduced the amount of hydroelectric power produced in the state. Such deficiency in usable water has increased the water prices in Los Angeles by 6-7% annually for the last decade and is expected to rise by 3.4% annually for typical customers for the next five years. For customers who use more than 20000 gallons a year, the bill is expected to increase by 34% by 2021. This price increase is occurring because of both drought and maintenance. <ref>Pacific Institute: The California Drought. 2014.</ref>

[[File:Figure_1.jpg|right|200px|frame|Figure 1. California Drought Index.]]
As of March 2015 the Metropolitan Water District of Southern California was willing to pay as much as $700 per acre-foot of water to farmers in the Sacramento Valley, so we assume this will be the target price of the water produced by the this desalination plant.<ref>Hessel, Phil. “Dry Southern California Offers Northern Farmers Top Dollar for Water”. NBC News. March 18, 2015. </ref>

Currently, there is a plant in Carlsbad which supplies drinking water to the San Diego area. The plant can produce up to 50 million gallons per day and is the nation's largest water desalination plant. The water from this plant costs $2257 per acre-foot for the first 48000 acre-feet and about $2000 for any additional acre-foot. The county signed a contract with this plant to purchase the first 48000 acre-feet of water for the next 30 years.<ref>Carlsbad Desalination Plant. “Project Overview”. January 2016.</ref> Also, there is another desalination plant in Tampa that produces about 25 million gallons of water per day, and over 2000 other public and industrial plants with production capacity of over 300000 gallons per day. When conducting the economic analysis our plant, such numbers presented by other plants must be considered, so our project can yield positive profit.<ref>Leven, Rachel. “U.S. Desalination Industry Grows Since 2000; See as Essential to Meeting Supply Needs”. Bloomberg BNA. August 21, 2013.</ref>

The profitability of the project is naturally heavily dependent on the price of water in California. If the drought ends, water supply will increase and the water price will drop significantly, but if the drought continues, the price of water in California could rise even more rapidly and allow desalination plants to be more profitable.<ref>Southern California Public Radio. “Where is California water use decreasing?”. December 2015.</ref> It may be difficult to make this plant profitable without a guaranteed purchase or other subsidy from the local governments because the presence of pre-existing plants maintains the water price relatively low. However, external sources of profit such as guaranteed purchase or subsidy are expected because the ramifications of not having desalination plant or other sources of water may be significantly more expensive than the subsidizing a desalination plant assuming the drought continues.<ref>Weiser, Matt. “Could desalination solve California’s water problem?”. The Sacramento Bee. October 18, 2014.</ref>

=Design Basis and Technical Approach=

[[File:Figure_2.jpg|left|200px|frame|Figure 2. Vertical Well Feed]]
Considering the number of pre-existing desalination plant and the potential risk of the drought ending, our plant will be tentatively designed to produce 15 million gallons of drinkable water per day to homes and businesses in the Los Angeles county area. With a recovery rate of near 30%, and a requirement for diluting any brine, the input feed flowrate will be 70.5 million gallons per day. Following the similarly built Carlsbad plant, we expect the plant to last 30 years. The project building timeline is one year based on our relatively small plant size. We will run for 350 days a year while shutting down for cleaning and maintenance once every 3 months (based on DOW standards).<ref name = "Bates">Bates, Wayne. “Cleaning your RO”. Hydranautics. </ref> The plant will be built 500 meters from the coastline on Long Beach. A borehole will be drilled 10 meters into the ground. At this depth there should be seawater aquifers below the coast that regularly recharges itself with seawater from the ocean. A vertical well containing intake holes will be put into the borehole. The intake holes will be equipped with large size filters that can easily be replaced. At the surface, the well will be connected to a large-scale pump in a well pump house. From there, the seawater will be horizontally pumped into a seawater tank in the desalination plant. This will be considered the start of our process. The process ends after pretreatment, desalination, post treatment, and brine treatment are completed. Appendix 1 shows more details on the Design Basis.

The feed composition is assumed to be the standard seawater “average” based on the DOW Chemical ROSA 9.0 Technical Manual in Table 1. However, in reality the seawater composition at the plant site will have to be chemically tested. The final product will be based on L.A. standards for drinking water quality (Table 1). The specifications in the table are upper limits that we cannot exceed. We will aim to reach the salt level standards but any remineralization and pH treatment will be exported and paid for due to the added complexity and cost in making such a system in-house.

Due to the acidity of the desalinated water coming out of the reverse osmosis (RO) system (pH ~ 6), the pumps and pipes carrying the desalinated water for post treatment will likely corrode over time. We have decided to place the desalination plant by the Long Beach Water Reclamation Plant (about 50 meters away) to prevent pipe corrosion.

[[File:Table_1.jpg|center|200px|frame|Table 1. Concentration of the standard seawater, the output from the RO system, and LA standard.]]

=Desalination Overview and Alternatives=

We first began our research into the current landscape of the desalination industry, particularly of plants that were built in the United States. We discovered there were many processes but the main ones included: 
* Membrane Separation (Forward Osmosis and Reverse Osmosis)
* Thermal Separation (Multi Stage Flash Distillation)
* Phase Change Separation (Freeze Separation)
Each process was analyzed for their pros and cons. For our project design, we decided to go with reverse osmosis as our desalination process. This is because it is the most commonly used process in the industry, and for good reason. Unlike thermal separation, which requires huge amounts of energy to supply heat, reverse osmosis is very energy efficient. This reduces energy costs that would be spread across the plant’s lifetime.<ref name = "Republic">Republic of Mauritius NY. "Technical Aspects of Desalination Plant". Sustainable Sanitation and Water Management. </ref>

Forward osmosis is a newer technology being developed that also uses membranes to separate water and salt and may be more energy efficient overall than reverse osmosis. This is because forward osmosis doesn’t require a pump to pressurize seawater through the membranes, but rather leaves a higher salt concentrated solution (AKA draw solution) on the other side of the membrane for the seawater to be drawn into. Then a second step is required to regenerate the concentrated draw solution and produce purified water.<ref>Mallinson, Alissa. "Study Shows Forward Osmosis Desalination Not Energy Efficient." MIT News [Boston]. July 23, 2014.</ref> We decided to move away from forward osmosis because it is a process that is still being refined and developed for desalination and lacks the data, programs, and technologies that reverse osmosis users benefit from. Additionally, a recent study from MIT has claimed that due to the inherent nature of forward osmosis requiring that seawater be drawn to a higher concentrated salt solution before the water and salt is separated, it cannot possibly be more energy efficient than reverse osmosis.<ref>O'Hern, Sean C., Karnik, Rohit. Nanofiltration across Defect-Sealed Nanoporous Monolayer Graphene. American Chemical Society 15.5 (2014); 3254-260. </ref>

Other processes such as freeze separation were too restrictive in their feed and process demands and required a lot of capital. Ultimately, reverse osmosis is simply the best process that exists for desalination and has been ranked the best when performance is evaluated in terms of technical, environmental, and economic aspects. It is a process that is flexible in water quantity, operation start-up and shut-off, and site location.

Nevertheless there are disadvantages to reverse osmosis that must be considered. The feed seawater must not be too salty and too low quality. The pressure to pump the seawater is enormous. The construction time is long for large plants. The capital costs are also high when compared with other processes such as thermal separation. But all of these disadvantages are ultimately outweighed by the advantages reverse osmosis provides.<ref name = "Republic"/>

=Process Overview and Alternatives=

See Appendix 2 (Figure 1) for the Process Flow Diagram (PFD). The process itself is detailed below. The stream flow rates are in Appendix 2 (Table 1). We assumed 10 meters of pipe for every stream except the inlet and outlet streams of the entire PFD.

==Pretreatment==

After the 10 meter vertical well pumps seawater from the Long Beach coast, the feed is pumped 500 meters to a seawater tank that makes it easy to shut-down the system and reduce the large volume of flow. The level control keeps the seawater from overfilling the tank. From there, part of the seawater goes to brine treatment while the rest go to pretreatment. The purpose of pretreatment in the desalination process is to prevent equipment damage by removing particles and microorganisms from the water. A variety of pollutants -- including calcium carbonate, sulfates, silica, clay, bacteria, and other colloids and microorganisms -- can disrupt the reverse osmosis process. Media filtration uses fine grains such as sand or anthracite to filter the particles. We used a large filter to catch larger particle material (e.g. sand). However, simple filtration is not enough to prevent colloid fouling. An ultrafilter (i.e. microfiltration) is used to remove virtually all suspended matter, effectively preventing colloid fouling. The costs associated with ultrafiltration are not prohibitive in terms of the project’s budget.<ref name = "Desalination">Desalination Pre-treatment. Lenntech.</ref>

Reverse osmosis systems can also be disrupted by the growth of microbiological films of bacteria on the film and pipe walls. They are typically treated as colloids during pretreatment, and most organisms are removed from the water by ultrafiltration, but this does not remove the smallest of organic particles. DBNPA is the most effective biocide, but these aren’t often used and have to be removed from the water. Chlorination is an effective way to remove bacteria, but it can cause degradation of the membrane, and must be removed before the reverse osmosis occurs. In our process we will use UV irradiation to remove any organic material that passes through the microfilters.<ref name = "Desalination"/>

Scaling can reduce the flux across the membrane and can build up in pipe corners, causing backup. It is caused by inorganic salts -- specifically sulfates and calcium and magnesium carbonate -- present in the water. Scaling can be prevented and/or controlled with acids, antiscalants, or ion exchange softening. Antiscalants are good for removing sulfates, but not much else. Acidification will only be effective at removing CaCO3 from water but, the decreased pH may deteriorate the membrane and pipes, so it may require future research.<ref name = "Desalination"/> Ion exchange is very expensive and requires more capital to undergo, but is overall the more effective option.

The ion exchanger is a pretreatment system which acts as a method to fight membrane scaling in reverse osmosis. The ion exchanger works by using sodium carboxylate as a resin which will react with Chlorine and Magnesium ions, removing them from the feed. The sodium is substituted out and enters the feed stream, increasing the overall sodium level and TDS. The lack of magnesium and chlorine ions means that there is nothing to react with the carbonates to form scaling salts on the membranes.

==Desalination (Reverse Osmosis)==

Two pumps (P-105 & P-106) take the adjusted feed seawater up to 58 bar. However, before this pumping occurs, about 70% of the original stream gets separated and heads into a PX Pressure Exchanger. This separated stream later joins together with the pumped stream to enter the reverse osmosis system at the required pressure of 58 bar. The streams can combine together as they have equal pressure. It is the concentrate stream from the output of the reverse osmosis system that transfers its energy to the adjusted feed in the PX Pressure Exchanger.

[[File:Figure_3.jpg|left|300px|frame|Figure 3. Inside look of pressure exchanger]]
The pressure of the concentrate stream is about 57 bar and it is fed into an exchanger which works by using vessel that resembles a small pipe within a rotating wheel. The low pressure (LP) feed enters the vessel at 1 bar (see Figure 3). The feed turns and is pushed out into an outlet pipe by the high pressure (HP) concentrate stream at ~57 bar. The feed is now a HP stream via this exchanged energy. The now LP concentrate stream turns and is pushed out into an outlet pipe by the LP feed entering the vessel.

98% of the energy within the concentrate stream can be recovered via the PX Pressure Exchanger.<ref>Desalination Product and Applications Catalog 2013. “Advance Technologies for Seawater Reverse Osmosis and Brackish Water Applications”. Energy Recovery, Inc. 2012-2014.</ref >Thus, the now HP feed stream is not quite 58 bar and will need to be pumped by a small booster pump (P-107).

When designing a reverse osmosis system, we need to make decisions on what specific kind of reverse osmosis system will be implemented. When making the decision, we investigated the advantages of different types of reverse osmosis among the following options: batch vs. continuous process, number of stages, and number of elements.<ref>FILMTEC Reverse Osmosis Membranes. “Water & Process Solutions Technical Manual.” Dow Chemical Company.</ref> Our decision was made based on process optimization which is described later in this report. The end result of that optimization was to go with a single stage, continuous process with five elements in series.

=Post Treatment=

The permeate is sent to post treatment after the salt content is analyzed and controlled if too high or too low due to any unexpected errors. As stated above, the post treatment process will be done 50 meters from our plant by the Long Beach Water Reclamation Plant. The reason we aren’t performing post treatment “in-house” is due to the increased cost it would take for us to build and run the equipment for post treatment. We have the option to use the government’s water treatment facilities that already exist and our 15 million gallons per day output is a relatively small addition to the large Long Beach Water Reclamation Plant. We will assume any costs to post treatment will be pushed to the consumer in the price of purchasing the drinking water from the Long Beach Water Reclamation Plant.

Post treatment at any facility works to bring water to “taste standards”. The water that is the permeate of our process is already meeting L.A. drinking standards as shown above in Table 1 after our mass balance calculations around the reverse osmosis system (Appendix 3, Table 1). The post treatment increases the pH from 6.3 to 7.3 with the addition of base and then adds minerals like fluoride, calcium, and magnesium back in to improve taste.

==Brine Treatment and Byproduct Waste==

A natural consequence of a reverse osmosis procedure is that a lot of salt will be leftover from the membrane separation in a very concentrated form. This is called “brine” and will have to be taken care of. According to EPA requirements, the salt concentration of water entering the environment (via land or sea) cannot be higher than 40,000 ppm.8 The salt concentration of our brine is much higher (around 70,000 ppm). Thus, the brine must be removed or treated. We decided to treat the brine by diluting it as removing brine can be difficult and may require lots of land (e.g. evaporation ponds).

After the concentrate exchanges energy in the energy exchanger, it becomes a low pressure stream. It is pumped into a dilution tank and mixes with 18.2 million gallons per day of seawater from the inlet to become diluted brine which is pumped 600 meters out to sea.

The waste byproducts are the solid buildup on the filters and membranes. The filters and membranes also have a set lifetime and thus will also need to be removed. The waste is all non-hazardous and thus can be put in the municipal dump.

=Process Details and Equipment Sizing=

==Filters and UV Treatment [Pretreatment]==
The cost of our ultrafiltration system is estimated using data from the University of Tennessee at Knoxville.<ref>Qiang He and R. Bruce Robinson Membrane Filtration Systems. “Membrane Module Costs.” University of Tennessee at Knoxville. December 2014.</ref> It estimated the cost of the cost of the ultrafiltration $0.057 per 1000 gallons and the construction and capital costs are $0.39 per gallon per day. The cost of UV treatment is based off of units available from buyultraviolet.com for units that can treat about 450 gallons per minute.<ref>Atlantic Ultraviolet Corporation. “Megatron Water Disinfection System.”</ref> Due to the water having to be in relatively small pipes for treatment we do not expect the costs to scale very much, so we believe these units to be an accurate cost estimate.

==Ion Exchanger [Pretreatment] ==

The ion exchanger was sized using data from DOW. Our process will use 60 vessels with ion exchange beds of volume 5.73 m3. With a flow rate of 30 bed volumes/hour (10,312.2 m3/hr) the system of beds has a capacity of 125% of the feed. This allows for periodic regeneration of the resin, a weak acid cation resin (DOW MAC-3 Resin). This weak cation resin will remove cationic calcium and magnesium from the water to prevent scaling. The mass flow was calculated via a mass balance on the ion exchanger (Appendix 3).

==Reverse Osmosis System [Desalination]==

Each of the 2607 Reverse Osmosis SW30HRLE-440i [Seawater 30 High Rejection Low Energy 440 ft2 interlocking cap membrane] units have the same dimensions. A, the net length, is 1016 mm. B, the total length including the endcap, is 1029 mm. C, the diameter of the filter element, is 201mm. D, the diameter of the permeate pathway, is 29 mm. For a single stage, we have five of the elements in series so the overall length of a single stage is expected to be roughly 5145 mm. The membrane material is polyamide thin film composite. It will last three years.
Operating Conditions [ROSA 9.0]: 
* Temperature: 30oC 
* Pressure: 58 bar 
* Flow Rate: 52.3 Million gal per day 
* Recovery: 28.7%

[[File:Figure_4.png|center|400px|frame|Figure 4. Diagram of Dow FILMTEC SW30HRLE-440i]]
Cleaning Schedule: 
* Conservative Estimate: Once every three months (three days available for cleaning)<ref name = "Bates"/> 
* DOW FILMTEC Cleaning Conditions 
** The normalized permeate flow drops 10%
* The normalized salt passage increases 5 - 10%

==Seawater Holding Tank and Dilution Tank [Brine Treatment]==

The initial seawater holding tank works as a buffer to prevent surges and help perform shutdowns quickly and easily. We expect to run continuously for 350 days a year with the rest of the available time spread out and used for shutting down the plant for cleaning membranes, filters, and the fixing of unexpected errors. There is a big tradeoff between having a holding tank with high retention time and having a holding tank with low cost. With that in mind, we erred on the side of lowering capital cost and went with a retention time of one hour. The volume could be found by multiplying the retention time with the feed flow rate (Stream 1). We assumed that a good height to width ratio for a large water tank would be 4:1.

A mass balance was done around the dilution tank to find what amount of seawater would be necessary to dilute the brine to 45,000 ppm salinity (Appendix 3; Table 3). With this the feed flow rate into the tank could be found. It was found that it takes about 20 minutes of retention time for good dilution of the brine in seawater.8 We again assumed that a good height to width ratio for a large water tank would be 4:1.

We assumed a corrosion allowance of 4 mm wall thickness. We calculated wall thickness using vessel pressure, hoop stress, longitudinal stress, and weld factor. Each tank is made with SS316 to protect against corrosion.<ref>Towler G, Sinnott R. Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013.</ref>

<center>
'''Table 2.''' Tank sizes.
[[File:Table_2.jpg|center|600px]]
</center>

=Process Optimization=
<center> '''Table 3.'''Optimization Criteria
[[File:Table_3.jpg|center|400px]]
</center>
In order to optimize the reverse osmosis system, numerous simulations were run. Three representative cases are reported as shown above (Table 3). The objective of the optimization was to minimize the required water price to break even in 10 years. In order to achieve this goal, the balance between the capital cost from the number of the filter elements and utility cost from energy consumption was weighed. Before the energy recovery system was implemented, the general trend was that if we have more filter elements in series, overall recovery increases but recovery rate per filter decreases, and we need more filter elements. While increasing the elements in series increase the required number of filters, and thereby increases the capital cost, we require less energy to produce the same amount of water. After the energy recovery system is implemented, the correlation between the number of elements in series and the required energy becomes more complicated because high pressure concentrate is used to pressurize the feed. However, the trade-off between the utility cost and capital cost still exists. Of the three cases shown above, single stage, five elements RO system turned out to be the most optimal solution because single stage, single element system required higher utility cost and two stage system slightly did not meet the LA standard. Detailed information which was used for the optimization is available in Appendix 3 and in Appendix 6.

=Ethical, Environmental, and Societal Considerations=

There are a few key considerations we have followed to meet ethical, environmental, and societal guidelines and requirements: 
* Meet L.A. Drinking Water Standards
** Correct ion concentration
* Moved near water treatment plant
** Pipes will not corrode from low pH
* Working below pressure limits
* Water will be locally served and meet taste standards
** Via Post Treatment mineralization
* Follows EPA Environmental Salinity Requirements
* Reduced Energy Requirements
** PX Pressure Exchanger
** Optimized RO System
** Reverse Osmosis is itself Energy Saving

For further environmental considerations, see effluent summary in Appendix 4.

=Economic Analysis=

The pricing of the majority of the process was completed using the Aspen Economic Analyzer, however we had determine costs for the ultrafiltration, reverse osmosis membranes, and the ultraviolet treatment was completed separately. This led to a total ISBL of $56 MM, OSBL of $22.4 MM, Engineering Costs of $22.4 MM and a Contingency of $7.8 MM. We also chose a working Capital of $10 MM. This working capital is sufficient because we have no raw materials costs. We also had FCOP of $5.2 MM and VCOP of $3.5 MM.

We originally were planning on pricing our water at $700 per acre-foot, but after doing economic analysis we had a 10 year NPV of $-76.2 MM. After determining did a sensitivity analysis and determined that if our water was priced at $2104 per acre-foot we could have an NPV of $0 after 10 years and if we were able to get the same price of $2257 per acre-foot as the plant in Carlsbad we would have a 10 year NPV of $8.3 MM which increases to $41.1 MM at 20 years. For more detailed data about the economic analysis see Appendix 6.

=Conclusion and Recommendations=

Our mission was to create an economically viable process that will desalinate water for human consumption in the L.A. County. We have found after research and optimization that the best process to undergo desalination is a continuous single stage, five elements reverse osmosis system. Our process can meet the required L.A. drinking water standard and we have set up plans for post treatment at a local water reclamation site. While meeting the drinking water standards, we also met our ethical, environmental, and social standards.

Nevertheless, at a sale price of $700 per acre-foot, our proposed desalination plant in Los Angeles, California looks to be a very expensive project to undergo. Without any sort of government assistance to carry the heavy capital costs of such a project, we would suggest the plant not be built. However, after performing a sensitivity analysis at higher (and less conservative) price points we reached a breakeven point in 10 years with a sale price of $2104 per acre-foot of drinking water. Research has shown that the droughts striking California look only to get worse in the near future. With limited freshwater resources, there is no option other than to consider desalination. Therefore it is reasonable to expect we could get near the same price deal as local desalination plants like the Carlsbad Plant ($2257 per acre-foot) As we have met all design constraints for this project, we hope to convince policymakers to adopt our project at least $2104 per acre-foot water. If this deal is reached, we recommend that this project and design be funded.

However, there are ways to improve our design with future work. Right now all the pipes diameters are based on having a flowrate of 1.5 m/s. But in reality, we’d need to use standard piping sizes to get more accurate pricing and energy costs. Also, in reality we would be purchasing some of these vessels in bulk quantities (i.e. bulk filters) so the price could be lowered with negotiation and economies of scale. We also have pumps that are providing very little pressure for their high cost. We can remove these pumps and simply run at higher pressures earlier. We can potentially reduce early energy costs by building a “floating” desalination plant in the sea to reduce pumping distance from the well.

=Appendices=

==Appendix 1: Design Basis==
[[File:Appendix_1.png|thumb|border|center|800px|middle|]]

==Appendix 2: Process Flow Diagram==
[[File:Appendix_2.png|center]]
==Appendix 3: Mass and Energy Balance Tables==
'''Table 1'''. Mass balance on 1 stage 5 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T1.png|center|500px|]]

'''Table 2'''. Mass balance on 1 stage 1 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T2.png|center|500px|]]

'''Table 3'''. Mass balance on 2 stage 5,5 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T3.png|center|600px|]]

'''Table 4'''. 1 stage 5 elements CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T4.png|center|400px|]]

'''Table 5'''. 1 stage 1 element CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T5.png|center|400px|]]

'''Table 6'''. 2 stages 5,5 elements CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T6.png|center|400px|]]

'''Table 7'''. Energy balance on 1 stage 5 elements RO system
[[File:Appendix_3_T7.png|center|700px|]]

==Appendix 4: Effluent Summary==
[[File:Appendix_4.png|center]]

==Appendix 5: Cost of Production==
[[File:Appendix_5.png|center]]

==Appendix 6: Economic Anlaysis==
===Case I ===
Water is $700; RO with 1 stage and 5 elements
[[File:Appendix_6_C1.png|center]]

===Case II===
Water is $1400; RO with 1 stage and 5 elements
[[File:Appendix_6_C2.png|center]]

===Case III: Carlsbad ===
Water is $2257; RO with 1 stage and 5 elements
[[File:Appendix_6_C3.png|center]]

===Case IV: Breakeven (NPV = 0 at 10 years)===
Water is $2104; RO with 1 stage and 5 elements
[[File:Appendix_6_C4.png|center]]

===Case V: Breakeven (NPV = 0 at 10 years)===
Water is $2112; RO with 2 stages and 5,5 elements
[[File:Appendix_6_C5.png|center]]

===Case VI: Breakeven (NPV = 0 at 10 years)===
Water is $2138 with 1 stage and 1 element
[[File:Appendix_6_C6.png|center]]

==Appendix 7: ROSA Simulation Outputs==

'''Table 1'''. System detail of 1 stage 5 elements RO system
[[File:Appendix_7_T1.png|center]]

'''Table 2'''. System detail of 1 stage 1 elements RO system
[[File:Appendix_7_T2.png|center]]

'''Table 3'''. System detail of 2 stage 5,5 elements system
[[File:Appendix_7_T3.png|center]]

'''Table 4'''. 1 stage 5 elements RO system concentrations and pH in each stream
[[File:Appendix_7_T4.png|center]]

'''Table 5'''. 1 stage 1 element RO system concentration and pH in each stream
[[File:Appendix_7_T5.png|center]]

'''Table 6'''. 2 stage 5,5 elements RO system concentrations and pH in each stream
[[File:Appendix_7_T6.png|center]]

=References=
<references/>

Desalination - Team E

2016-03-12T03:33:06Z

Bts511: /* Appendix 3: Mass and Energy Balance Tables */

Group E Corporation

Authors: Hassan Ali, Woo Soo Choe, Brett Sleyster, Jake Stolley

Instructors: Fengqi You, David Wegerer

March 11, 2016

=Executive Summary=

Outlined in this report is the project model report for the proposed desalination plant in Los Angeles, California which will provide drinking water to local citizens using reverse osmosis process. This location was chosen due to the region’s propensity for droughts in recent years. The fresh water resources in California have plummeted as water prices have risen 3-4% each year. This is only expected to continue. Other companies have worked deals with the Californian government to create desalination plants that provide a new fresh water drinking source. This suggests that there is a strong market here with potential growth.

The plant will have an output of 15 million gallons per day, which is about medium size relative to other plants. Thus the build time will be about one year. The plant will run for 30 years, with 350 days per year of use. Shutdown time will be used for cleaning and maintenance of equipment. The plant will be built on Long Beach, 500 meters from the coast nearby the Long Beach Water Reclamation Plant (LBWP). The input will withdraw 70.5 million gallons per day of seawater that will be used for reverse osmosis and dilution of brine.

Reverse osmosis was chosen as the desalination method for its low energy cost and flexibility. We designed a process that will pretreat the seawater to remove waste, colloids, bacteria, and membrane-damaging scalants. Once this seawater is treated is pumped to heavy pressures through a reverse osmosis system. After doing process optimization under various system parameter changes, we found a continuous, single stage, five element reverse osmosis process to be best. It produces a near 30% recovery. We reduce energy costs by including a pressure exchanger between our concentrate outlet and adjusted feed into the reverse osmosis system.

After performing mass and energy balances, we were able to size equipment and find our permeate composition. It turns out our process will meet L.A. standards for drinking water quality. However, post treatment will be done by the local LBWP to make the water taste better and be non-caustic to pipes. This will reduce “in-house” costs. Many environmental considerations were made, to meet our environmental standards. We meet EPA salinity requirements for our brine waste by diluting it. We also don’t use a lot of energy and thus reduce our fossil fuel consumption and CO2 generation. To be safe, we work well under max pressure limits for our equipment. We make sure we are ethical and meet all standards that our local county requires of us.

After an economic evaluation at a conservative well-water price of $700 per acre-foot, we realize that our revenue is way too low to cover our costs. Capital costs are especially large with regards to our large water tanks. We find that we reach a 10 year NPV breakpoint of $2104 per acre-foot after performing a sensitivity analysis. The Carlsbad Plant in San Diego worked out a deal for $2257 per acre-foot.

As we have met all design constraints for this project, we hope to convince policymakers to adopt our project at least $2104 per acre-foot water. If this deal is reached, we recommend that this project and design be funded. There is future work to be done to improve costs such as performing pump optimization. However, this work is minimal and won’t harm our bottom line if we reach our proposed price points.

__TOC__

=Introduction=

Due to a combination of various factors--including rising human population, climate change, increased energy consumption, and land erosion--the availability of fresh water as a resource for human consumption has been dwindling. This has had a negative impact on the world, but this impact has been especially felt in areas with natural shortages in fresh water. Desert regions, such as Saudi Arabia, have resorted to desalinating seawater at the coast and transporting up to their cities. This encompasses their entire freshwater system as very few natural water sources remain. In the United States there is a region which has shown increasing signs of following in the footsteps of Saudi Arabia and becoming fully reliant on desalination: Los Angeles, California.

=Location and Market Analysis=

In this project, Los Angeles was chosen as the location our plant, because the area has been subject to a severe drought for the last four years, as shown in Figure 1. The rapid decrease in the available water has caused permanent damages to the ground, decreased the amount of water in the reservoir from 45% to 25%, and reduced the amount of hydroelectric power produced in the state. Such deficiency in usable water has increased the water prices in Los Angeles by 6-7% annually for the last decade and is expected to rise by 3.4% annually for typical customers for the next five years. For customers who use more than 20000 gallons a year, the bill is expected to increase by 34% by 2021. This price increase is occurring because of both drought and maintenance. <ref>Pacific Institute: The California Drought. 2014.</ref>

[[File:Figure_1.jpg|right|200px|frame|Figure 1. California Drought Index.]]
As of March 2015 the Metropolitan Water District of Southern California was willing to pay as much as $700 per acre-foot of water to farmers in the Sacramento Valley, so we assume this will be the target price of the water produced by the this desalination plant.<ref>Hessel, Phil. “Dry Southern California Offers Northern Farmers Top Dollar for Water”. NBC News. March 18, 2015. </ref>

Currently, there is a plant in Carlsbad which supplies drinking water to the San Diego area. The plant can produce up to 50 million gallons per day and is the nation's largest water desalination plant. The water from this plant costs $2257 per acre-foot for the first 48000 acre-feet and about $2000 for any additional acre-foot. The county signed a contract with this plant to purchase the first 48000 acre-feet of water for the next 30 years.<ref>Carlsbad Desalination Plant. “Project Overview”. January 2016.</ref> Also, there is another desalination plant in Tampa that produces about 25 million gallons of water per day, and over 2000 other public and industrial plants with production capacity of over 300000 gallons per day. When conducting the economic analysis our plant, such numbers presented by other plants must be considered, so our project can yield positive profit.<ref>Leven, Rachel. “U.S. Desalination Industry Grows Since 2000; See as Essential to Meeting Supply Needs”. Bloomberg BNA. August 21, 2013.</ref>

The profitability of the project is naturally heavily dependent on the price of water in California. If the drought ends, water supply will increase and the water price will drop significantly, but if the drought continues, the price of water in California could rise even more rapidly and allow desalination plants to be more profitable.<ref>Southern California Public Radio. “Where is California water use decreasing?”. December 2015.</ref> It may be difficult to make this plant profitable without a guaranteed purchase or other subsidy from the local governments because the presence of pre-existing plants maintains the water price relatively low. However, external sources of profit such as guaranteed purchase or subsidy are expected because the ramifications of not having desalination plant or other sources of water may be significantly more expensive than the subsidizing a desalination plant assuming the drought continues.<ref>Weiser, Matt. “Could desalination solve California’s water problem?”. The Sacramento Bee. October 18, 2014.</ref>

=Design Basis and Technical Approach=

[[File:Figure_2.jpg|left|200px|frame|Figure 2. Vertical Well Feed]]
Considering the number of pre-existing desalination plant and the potential risk of the drought ending, our plant will be tentatively designed to produce 15 million gallons of drinkable water per day to homes and businesses in the Los Angeles county area. With a recovery rate of near 30%, and a requirement for diluting any brine, the input feed flowrate will be 70.5 million gallons per day. Following the similarly built Carlsbad plant, we expect the plant to last 30 years. The project building timeline is one year based on our relatively small plant size. We will run for 350 days a year while shutting down for cleaning and maintenance once every 3 months (based on DOW standards).<ref name = "Bates">Bates, Wayne. “Cleaning your RO”. Hydranautics. </ref> The plant will be built 500 meters from the coastline on Long Beach. A borehole will be drilled 10 meters into the ground. At this depth there should be seawater aquifers below the coast that regularly recharges itself with seawater from the ocean. A vertical well containing intake holes will be put into the borehole. The intake holes will be equipped with large size filters that can easily be replaced. At the surface, the well will be connected to a large-scale pump in a well pump house. From there, the seawater will be horizontally pumped into a seawater tank in the desalination plant. This will be considered the start of our process. The process ends after pretreatment, desalination, post treatment, and brine treatment are completed. Appendix 1 shows more details on the Design Basis.

The feed composition is assumed to be the standard seawater “average” based on the DOW Chemical ROSA 9.0 Technical Manual in Table 1. However, in reality the seawater composition at the plant site will have to be chemically tested. The final product will be based on L.A. standards for drinking water quality (Table 1). The specifications in the table are upper limits that we cannot exceed. We will aim to reach the salt level standards but any remineralization and pH treatment will be exported and paid for due to the added complexity and cost in making such a system in-house.

Due to the acidity of the desalinated water coming out of the reverse osmosis (RO) system (pH ~ 6), the pumps and pipes carrying the desalinated water for post treatment will likely corrode over time. We have decided to place the desalination plant by the Long Beach Water Reclamation Plant (about 50 meters away) to prevent pipe corrosion.

[[File:Table_1.jpg|center|200px|frame|Table 1. Concentration of the standard seawater, the output from the RO system, and LA standard.]]

=Desalination Overview and Alternatives=

We first began our research into the current landscape of the desalination industry, particularly of plants that were built in the United States. We discovered there were many processes but the main ones included: 
* Membrane Separation (Forward Osmosis and Reverse Osmosis)
* Thermal Separation (Multi Stage Flash Distillation)
* Phase Change Separation (Freeze Separation)
Each process was analyzed for their pros and cons. For our project design, we decided to go with reverse osmosis as our desalination process. This is because it is the most commonly used process in the industry, and for good reason. Unlike thermal separation, which requires huge amounts of energy to supply heat, reverse osmosis is very energy efficient. This reduces energy costs that would be spread across the plant’s lifetime.<ref name = "Republic">Republic of Mauritius NY. "Technical Aspects of Desalination Plant". Sustainable Sanitation and Water Management. </ref>

Forward osmosis is a newer technology being developed that also uses membranes to separate water and salt and may be more energy efficient overall than reverse osmosis. This is because forward osmosis doesn’t require a pump to pressurize seawater through the membranes, but rather leaves a higher salt concentrated solution (AKA draw solution) on the other side of the membrane for the seawater to be drawn into. Then a second step is required to regenerate the concentrated draw solution and produce purified water.<ref>Mallinson, Alissa. "Study Shows Forward Osmosis Desalination Not Energy Efficient." MIT News [Boston]. July 23, 2014.</ref> We decided to move away from forward osmosis because it is a process that is still being refined and developed for desalination and lacks the data, programs, and technologies that reverse osmosis users benefit from. Additionally, a recent study from MIT has claimed that due to the inherent nature of forward osmosis requiring that seawater be drawn to a higher concentrated salt solution before the water and salt is separated, it cannot possibly be more energy efficient than reverse osmosis.<ref>O'Hern, Sean C., Karnik, Rohit. Nanofiltration across Defect-Sealed Nanoporous Monolayer Graphene. American Chemical Society 15.5 (2014); 3254-260. </ref>

Other processes such as freeze separation were too restrictive in their feed and process demands and required a lot of capital. Ultimately, reverse osmosis is simply the best process that exists for desalination and has been ranked the best when performance is evaluated in terms of technical, environmental, and economic aspects. It is a process that is flexible in water quantity, operation start-up and shut-off, and site location.

Nevertheless there are disadvantages to reverse osmosis that must be considered. The feed seawater must not be too salty and too low quality. The pressure to pump the seawater is enormous. The construction time is long for large plants. The capital costs are also high when compared with other processes such as thermal separation. But all of these disadvantages are ultimately outweighed by the advantages reverse osmosis provides.<ref name = "Republic"/>

=Process Overview and Alternatives=

See Appendix 2 (Figure 1) for the Process Flow Diagram (PFD). The process itself is detailed below. The stream flow rates are in Appendix 2 (Table 1). We assumed 10 meters of pipe for every stream except the inlet and outlet streams of the entire PFD.

==Pretreatment==

After the 10 meter vertical well pumps seawater from the Long Beach coast, the feed is pumped 500 meters to a seawater tank that makes it easy to shut-down the system and reduce the large volume of flow. The level control keeps the seawater from overfilling the tank. From there, part of the seawater goes to brine treatment while the rest go to pretreatment. The purpose of pretreatment in the desalination process is to prevent equipment damage by removing particles and microorganisms from the water. A variety of pollutants -- including calcium carbonate, sulfates, silica, clay, bacteria, and other colloids and microorganisms -- can disrupt the reverse osmosis process. Media filtration uses fine grains such as sand or anthracite to filter the particles. We used a large filter to catch larger particle material (e.g. sand). However, simple filtration is not enough to prevent colloid fouling. An ultrafilter (i.e. microfiltration) is used to remove virtually all suspended matter, effectively preventing colloid fouling. The costs associated with ultrafiltration are not prohibitive in terms of the project’s budget.<ref name = "Desalination">Desalination Pre-treatment. Lenntech.</ref>

Reverse osmosis systems can also be disrupted by the growth of microbiological films of bacteria on the film and pipe walls. They are typically treated as colloids during pretreatment, and most organisms are removed from the water by ultrafiltration, but this does not remove the smallest of organic particles. DBNPA is the most effective biocide, but these aren’t often used and have to be removed from the water. Chlorination is an effective way to remove bacteria, but it can cause degradation of the membrane, and must be removed before the reverse osmosis occurs. In our process we will use UV irradiation to remove any organic material that passes through the microfilters.<ref name = "Desalination"/>

Scaling can reduce the flux across the membrane and can build up in pipe corners, causing backup. It is caused by inorganic salts -- specifically sulfates and calcium and magnesium carbonate -- present in the water. Scaling can be prevented and/or controlled with acids, antiscalants, or ion exchange softening. Antiscalants are good for removing sulfates, but not much else. Acidification will only be effective at removing CaCO3 from water but, the decreased pH may deteriorate the membrane and pipes, so it may require future research.<ref name = "Desalination"/> Ion exchange is very expensive and requires more capital to undergo, but is overall the more effective option.

The ion exchanger is a pretreatment system which acts as a method to fight membrane scaling in reverse osmosis. The ion exchanger works by using sodium carboxylate as a resin which will react with Chlorine and Magnesium ions, removing them from the feed. The sodium is substituted out and enters the feed stream, increasing the overall sodium level and TDS. The lack of magnesium and chlorine ions means that there is nothing to react with the carbonates to form scaling salts on the membranes.

==Desalination (Reverse Osmosis)==

Two pumps (P-105 & P-106) take the adjusted feed seawater up to 58 bar. However, before this pumping occurs, about 70% of the original stream gets separated and heads into a PX Pressure Exchanger. This separated stream later joins together with the pumped stream to enter the reverse osmosis system at the required pressure of 58 bar. The streams can combine together as they have equal pressure. It is the concentrate stream from the output of the reverse osmosis system that transfers its energy to the adjusted feed in the PX Pressure Exchanger.

[[File:Figure_3.jpg|left|300px|frame|Figure 3. Inside look of pressure exchanger]]
The pressure of the concentrate stream is about 57 bar and it is fed into an exchanger which works by using vessel that resembles a small pipe within a rotating wheel. The low pressure (LP) feed enters the vessel at 1 bar (see Figure 3). The feed turns and is pushed out into an outlet pipe by the high pressure (HP) concentrate stream at ~57 bar. The feed is now a HP stream via this exchanged energy. The now LP concentrate stream turns and is pushed out into an outlet pipe by the LP feed entering the vessel.

98% of the energy within the concentrate stream can be recovered via the PX Pressure Exchanger.<ref>Desalination Product and Applications Catalog 2013. “Advance Technologies for Seawater Reverse Osmosis and Brackish Water Applications”. Energy Recovery, Inc. 2012-2014.</ref >Thus, the now HP feed stream is not quite 58 bar and will need to be pumped by a small booster pump (P-107).

When designing a reverse osmosis system, we need to make decisions on what specific kind of reverse osmosis system will be implemented. When making the decision, we investigated the advantages of different types of reverse osmosis among the following options: batch vs. continuous process, number of stages, and number of elements.<ref>FILMTEC Reverse Osmosis Membranes. “Water & Process Solutions Technical Manual.” Dow Chemical Company.</ref> Our decision was made based on process optimization which is described later in this report. The end result of that optimization was to go with a single stage, continuous process with five elements in series.

=Post Treatment=

The permeate is sent to post treatment after the salt content is analyzed and controlled if too high or too low due to any unexpected errors. As stated above, the post treatment process will be done 50 meters from our plant by the Long Beach Water Reclamation Plant. The reason we aren’t performing post treatment “in-house” is due to the increased cost it would take for us to build and run the equipment for post treatment. We have the option to use the government’s water treatment facilities that already exist and our 15 million gallons per day output is a relatively small addition to the large Long Beach Water Reclamation Plant. We will assume any costs to post treatment will be pushed to the consumer in the price of purchasing the drinking water from the Long Beach Water Reclamation Plant.

Post treatment at any facility works to bring water to “taste standards”. The water that is the permeate of our process is already meeting L.A. drinking standards as shown above in Table 1 after our mass balance calculations around the reverse osmosis system (Appendix 3, Table 1). The post treatment increases the pH from 6.3 to 7.3 with the addition of base and then adds minerals like fluoride, calcium, and magnesium back in to improve taste.

==Brine Treatment and Byproduct Waste==

A natural consequence of a reverse osmosis procedure is that a lot of salt will be leftover from the membrane separation in a very concentrated form. This is called “brine” and will have to be taken care of. According to EPA requirements, the salt concentration of water entering the environment (via land or sea) cannot be higher than 40,000 ppm.8 The salt concentration of our brine is much higher (around 70,000 ppm). Thus, the brine must be removed or treated. We decided to treat the brine by diluting it as removing brine can be difficult and may require lots of land (e.g. evaporation ponds).

After the concentrate exchanges energy in the energy exchanger, it becomes a low pressure stream. It is pumped into a dilution tank and mixes with 18.2 million gallons per day of seawater from the inlet to become diluted brine which is pumped 600 meters out to sea.

The waste byproducts are the solid buildup on the filters and membranes. The filters and membranes also have a set lifetime and thus will also need to be removed. The waste is all non-hazardous and thus can be put in the municipal dump.

=Process Details and Equipment Sizing=

==Filters and UV Treatment [Pretreatment]==
The cost of our ultrafiltration system is estimated using data from the University of Tennessee at Knoxville.<ref>Qiang He and R. Bruce Robinson Membrane Filtration Systems. “Membrane Module Costs.” University of Tennessee at Knoxville. December 2014.</ref> It estimated the cost of the cost of the ultrafiltration $0.057 per 1000 gallons and the construction and capital costs are $0.39 per gallon per day. The cost of UV treatment is based off of units available from buyultraviolet.com for units that can treat about 450 gallons per minute.<ref>Atlantic Ultraviolet Corporation. “Megatron Water Disinfection System.”</ref> Due to the water having to be in relatively small pipes for treatment we do not expect the costs to scale very much, so we believe these units to be an accurate cost estimate.

==Ion Exchanger [Pretreatment] ==

The ion exchanger was sized using data from DOW. Our process will use 60 vessels with ion exchange beds of volume 5.73 m3. With a flow rate of 30 bed volumes/hour (10,312.2 m3/hr) the system of beds has a capacity of 125% of the feed. This allows for periodic regeneration of the resin, a weak acid cation resin (DOW MAC-3 Resin). This weak cation resin will remove cationic calcium and magnesium from the water to prevent scaling. The mass flow was calculated via a mass balance on the ion exchanger (Appendix 3).

==Reverse Osmosis System [Desalination]==

Each of the 2607 Reverse Osmosis SW30HRLE-440i [Seawater 30 High Rejection Low Energy 440 ft2 interlocking cap membrane] units have the same dimensions. A, the net length, is 1016 mm. B, the total length including the endcap, is 1029 mm. C, the diameter of the filter element, is 201mm. D, the diameter of the permeate pathway, is 29 mm. For a single stage, we have five of the elements in series so the overall length of a single stage is expected to be roughly 5145 mm. The membrane material is polyamide thin film composite. It will last three years.
Operating Conditions [ROSA 9.0]: 
* Temperature: 30oC 
* Pressure: 58 bar 
* Flow Rate: 52.3 Million gal per day 
* Recovery: 28.7%

[[File:Figure_4.png|center|400px|frame|Figure 4. Diagram of Dow FILMTEC SW30HRLE-440i]]
Cleaning Schedule: 
* Conservative Estimate: Once every three months (three days available for cleaning)<ref name = "Bates"/> 
* DOW FILMTEC Cleaning Conditions 
** The normalized permeate flow drops 10%
* The normalized salt passage increases 5 - 10%

==Seawater Holding Tank and Dilution Tank [Brine Treatment]==

The initial seawater holding tank works as a buffer to prevent surges and help perform shutdowns quickly and easily. We expect to run continuously for 350 days a year with the rest of the available time spread out and used for shutting down the plant for cleaning membranes, filters, and the fixing of unexpected errors. There is a big tradeoff between having a holding tank with high retention time and having a holding tank with low cost. With that in mind, we erred on the side of lowering capital cost and went with a retention time of one hour. The volume could be found by multiplying the retention time with the feed flow rate (Stream 1). We assumed that a good height to width ratio for a large water tank would be 4:1.

A mass balance was done around the dilution tank to find what amount of seawater would be necessary to dilute the brine to 45,000 ppm salinity (Appendix 3; Table 3). With this the feed flow rate into the tank could be found. It was found that it takes about 20 minutes of retention time for good dilution of the brine in seawater.8 We again assumed that a good height to width ratio for a large water tank would be 4:1.

We assumed a corrosion allowance of 4 mm wall thickness. We calculated wall thickness using vessel pressure, hoop stress, longitudinal stress, and weld factor. Each tank is made with SS316 to protect against corrosion.<ref>Towler G, Sinnott R. Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013.</ref>

<center>
'''Table 2.''' Tank sizes.
[[File:Table_2.jpg|center|600px]]
</center>

=Process Optimization=
<center> '''Table 3.'''Optimization Criteria
[[File:Table_3.jpg|center|400px]]
</center>
In order to optimize the reverse osmosis system, numerous simulations were run. Three representative cases are reported as shown above (Table 3). The objective of the optimization was to minimize the required water price to break even in 10 years. In order to achieve this goal, the balance between the capital cost from the number of the filter elements and utility cost from energy consumption was weighed. Before the energy recovery system was implemented, the general trend was that if we have more filter elements in series, overall recovery increases but recovery rate per filter decreases, and we need more filter elements. While increasing the elements in series increase the required number of filters, and thereby increases the capital cost, we require less energy to produce the same amount of water. After the energy recovery system is implemented, the correlation between the number of elements in series and the required energy becomes more complicated because high pressure concentrate is used to pressurize the feed. However, the trade-off between the utility cost and capital cost still exists. Of the three cases shown above, single stage, five elements RO system turned out to be the most optimal solution because single stage, single element system required higher utility cost and two stage system slightly did not meet the LA standard. Detailed information which was used for the optimization is available in Appendix 3 and in Appendix 6.

=Ethical, Environmental, and Societal Considerations=

There are a few key considerations we have followed to meet ethical, environmental, and societal guidelines and requirements: 
* Meet L.A. Drinking Water Standards
** Correct ion concentration
* Moved near water treatment plant
** Pipes will not corrode from low pH
* Working below pressure limits
* Water will be locally served and meet taste standards
** Via Post Treatment mineralization
* Follows EPA Environmental Salinity Requirements
* Reduced Energy Requirements
** PX Pressure Exchanger
** Optimized RO System
** Reverse Osmosis is itself Energy Saving

For further environmental considerations, see effluent summary in Appendix 4.

=Economic Analysis=

The pricing of the majority of the process was completed using the Aspen Economic Analyzer, however we had determine costs for the ultrafiltration, reverse osmosis membranes, and the ultraviolet treatment was completed separately. This led to a total ISBL of $56 MM, OSBL of $22.4 MM, Engineering Costs of $22.4 MM and a Contingency of $7.8 MM. We also chose a working Capital of $10 MM. This working capital is sufficient because we have no raw materials costs. We also had FCOP of $5.2 MM and VCOP of $3.5 MM.

We originally were planning on pricing our water at $700 per acre-foot, but after doing economic analysis we had a 10 year NPV of $-76.2 MM. After determining did a sensitivity analysis and determined that if our water was priced at $2104 per acre-foot we could have an NPV of $0 after 10 years and if we were able to get the same price of $2257 per acre-foot as the plant in Carlsbad we would have a 10 year NPV of $8.3 MM which increases to $41.1 MM at 20 years. For more detailed data about the economic analysis see Appendix 6.

=Conclusion and Recommendations=

Our mission was to create an economically viable process that will desalinate water for human consumption in the L.A. County. We have found after research and optimization that the best process to undergo desalination is a continuous single stage, five elements reverse osmosis system. Our process can meet the required L.A. drinking water standard and we have set up plans for post treatment at a local water reclamation site. While meeting the drinking water standards, we also met our ethical, environmental, and social standards.

Nevertheless, at a sale price of $700 per acre-foot, our proposed desalination plant in Los Angeles, California looks to be a very expensive project to undergo. Without any sort of government assistance to carry the heavy capital costs of such a project, we would suggest the plant not be built. However, after performing a sensitivity analysis at higher (and less conservative) price points we reached a breakeven point in 10 years with a sale price of $2104 per acre-foot of drinking water. Research has shown that the droughts striking California look only to get worse in the near future. With limited freshwater resources, there is no option other than to consider desalination. Therefore it is reasonable to expect we could get near the same price deal as local desalination plants like the Carlsbad Plant ($2257 per acre-foot) As we have met all design constraints for this project, we hope to convince policymakers to adopt our project at least $2104 per acre-foot water. If this deal is reached, we recommend that this project and design be funded.

However, there are ways to improve our design with future work. Right now all the pipes diameters are based on having a flowrate of 1.5 m/s. But in reality, we’d need to use standard piping sizes to get more accurate pricing and energy costs. Also, in reality we would be purchasing some of these vessels in bulk quantities (i.e. bulk filters) so the price could be lowered with negotiation and economies of scale. We also have pumps that are providing very little pressure for their high cost. We can remove these pumps and simply run at higher pressures earlier. We can potentially reduce early energy costs by building a “floating” desalination plant in the sea to reduce pumping distance from the well.

=Appendices=

==Appendix 1: Design Basis==
[[File:Appendix_1.png|thumb|border|center|800px|middle|]]

==Appendix 2: Process Flow Diagram==
[[File:Appendix_2.png|center]]
==Appendix 3: Mass and Energy Balance Tables==
'''Table 1'''. Mass balance on 1 stage 5 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T1.png|center|600px|]]

'''Table 2'''. Mass balance on 1 stage 1 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T2.png|center|600px|]]

'''Table 3'''. Mass balance on 2 stage 5,5 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T3.png|center|600px|]]

'''Table 4'''. 1 stage 5 elements CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T4.png|center|400px|]]

'''Table 5'''. 1 stage 1 element CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T5.png|center|400px|]]

'''Table 6'''. 2 stages 5,5 elements CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T6.png|center|400px|]]

'''Table 7'''. Energy balance on 1 stage 5 elements RO system
[[File:Appendix_3_T7.png|center|600px|]]

==Appendix 4: Effluent Summary==
[[File:Appendix_4.png|center]]

==Appendix 5: Cost of Production==
[[File:Appendix_5.png|center]]

==Appendix 6: Economic Anlaysis==
===Case I ===
Water is $700; RO with 1 stage and 5 elements
[[File:Appendix_6_C1.png|center]]

===Case II===
Water is $1400; RO with 1 stage and 5 elements
[[File:Appendix_6_C2.png|center]]

===Case III: Carlsbad ===
Water is $2257; RO with 1 stage and 5 elements
[[File:Appendix_6_C3.png|center]]

===Case IV: Breakeven (NPV = 0 at 10 years)===
Water is $2104; RO with 1 stage and 5 elements
[[File:Appendix_6_C4.png|center]]

===Case V: Breakeven (NPV = 0 at 10 years)===
Water is $2112; RO with 2 stages and 5,5 elements
[[File:Appendix_6_C5.png|center]]

===Case VI: Breakeven (NPV = 0 at 10 years)===
Water is $2138 with 1 stage and 1 element
[[File:Appendix_6_C6.png|center]]

==Appendix 7: ROSA Simulation Outputs==

'''Table 1'''. System detail of 1 stage 5 elements RO system
[[File:Appendix_7_T1.png|center]]

'''Table 2'''. System detail of 1 stage 1 elements RO system
[[File:Appendix_7_T2.png|center]]

'''Table 3'''. System detail of 2 stage 5,5 elements system
[[File:Appendix_7_T3.png|center]]

'''Table 4'''. 1 stage 5 elements RO system concentrations and pH in each stream
[[File:Appendix_7_T4.png|center]]

'''Table 5'''. 1 stage 1 element RO system concentration and pH in each stream
[[File:Appendix_7_T5.png|center]]

'''Table 6'''. 2 stage 5,5 elements RO system concentrations and pH in each stream
[[File:Appendix_7_T6.png|center]]

=References=
<references/>

Desalination - Team E

2016-03-12T03:32:48Z

Bts511: /* Appendix 3: Mass and Energy Balance Tables */

Group E Corporation

Authors: Hassan Ali, Woo Soo Choe, Brett Sleyster, Jake Stolley

Instructors: Fengqi You, David Wegerer

March 11, 2016

=Executive Summary=

Outlined in this report is the project model report for the proposed desalination plant in Los Angeles, California which will provide drinking water to local citizens using reverse osmosis process. This location was chosen due to the region’s propensity for droughts in recent years. The fresh water resources in California have plummeted as water prices have risen 3-4% each year. This is only expected to continue. Other companies have worked deals with the Californian government to create desalination plants that provide a new fresh water drinking source. This suggests that there is a strong market here with potential growth.

The plant will have an output of 15 million gallons per day, which is about medium size relative to other plants. Thus the build time will be about one year. The plant will run for 30 years, with 350 days per year of use. Shutdown time will be used for cleaning and maintenance of equipment. The plant will be built on Long Beach, 500 meters from the coast nearby the Long Beach Water Reclamation Plant (LBWP). The input will withdraw 70.5 million gallons per day of seawater that will be used for reverse osmosis and dilution of brine.

Reverse osmosis was chosen as the desalination method for its low energy cost and flexibility. We designed a process that will pretreat the seawater to remove waste, colloids, bacteria, and membrane-damaging scalants. Once this seawater is treated is pumped to heavy pressures through a reverse osmosis system. After doing process optimization under various system parameter changes, we found a continuous, single stage, five element reverse osmosis process to be best. It produces a near 30% recovery. We reduce energy costs by including a pressure exchanger between our concentrate outlet and adjusted feed into the reverse osmosis system.

After performing mass and energy balances, we were able to size equipment and find our permeate composition. It turns out our process will meet L.A. standards for drinking water quality. However, post treatment will be done by the local LBWP to make the water taste better and be non-caustic to pipes. This will reduce “in-house” costs. Many environmental considerations were made, to meet our environmental standards. We meet EPA salinity requirements for our brine waste by diluting it. We also don’t use a lot of energy and thus reduce our fossil fuel consumption and CO2 generation. To be safe, we work well under max pressure limits for our equipment. We make sure we are ethical and meet all standards that our local county requires of us.

After an economic evaluation at a conservative well-water price of $700 per acre-foot, we realize that our revenue is way too low to cover our costs. Capital costs are especially large with regards to our large water tanks. We find that we reach a 10 year NPV breakpoint of $2104 per acre-foot after performing a sensitivity analysis. The Carlsbad Plant in San Diego worked out a deal for $2257 per acre-foot.

As we have met all design constraints for this project, we hope to convince policymakers to adopt our project at least $2104 per acre-foot water. If this deal is reached, we recommend that this project and design be funded. There is future work to be done to improve costs such as performing pump optimization. However, this work is minimal and won’t harm our bottom line if we reach our proposed price points.

__TOC__

=Introduction=

Due to a combination of various factors--including rising human population, climate change, increased energy consumption, and land erosion--the availability of fresh water as a resource for human consumption has been dwindling. This has had a negative impact on the world, but this impact has been especially felt in areas with natural shortages in fresh water. Desert regions, such as Saudi Arabia, have resorted to desalinating seawater at the coast and transporting up to their cities. This encompasses their entire freshwater system as very few natural water sources remain. In the United States there is a region which has shown increasing signs of following in the footsteps of Saudi Arabia and becoming fully reliant on desalination: Los Angeles, California.

=Location and Market Analysis=

In this project, Los Angeles was chosen as the location our plant, because the area has been subject to a severe drought for the last four years, as shown in Figure 1. The rapid decrease in the available water has caused permanent damages to the ground, decreased the amount of water in the reservoir from 45% to 25%, and reduced the amount of hydroelectric power produced in the state. Such deficiency in usable water has increased the water prices in Los Angeles by 6-7% annually for the last decade and is expected to rise by 3.4% annually for typical customers for the next five years. For customers who use more than 20000 gallons a year, the bill is expected to increase by 34% by 2021. This price increase is occurring because of both drought and maintenance. <ref>Pacific Institute: The California Drought. 2014.</ref>

[[File:Figure_1.jpg|right|200px|frame|Figure 1. California Drought Index.]]
As of March 2015 the Metropolitan Water District of Southern California was willing to pay as much as $700 per acre-foot of water to farmers in the Sacramento Valley, so we assume this will be the target price of the water produced by the this desalination plant.<ref>Hessel, Phil. “Dry Southern California Offers Northern Farmers Top Dollar for Water”. NBC News. March 18, 2015. </ref>

Currently, there is a plant in Carlsbad which supplies drinking water to the San Diego area. The plant can produce up to 50 million gallons per day and is the nation's largest water desalination plant. The water from this plant costs $2257 per acre-foot for the first 48000 acre-feet and about $2000 for any additional acre-foot. The county signed a contract with this plant to purchase the first 48000 acre-feet of water for the next 30 years.<ref>Carlsbad Desalination Plant. “Project Overview”. January 2016.</ref> Also, there is another desalination plant in Tampa that produces about 25 million gallons of water per day, and over 2000 other public and industrial plants with production capacity of over 300000 gallons per day. When conducting the economic analysis our plant, such numbers presented by other plants must be considered, so our project can yield positive profit.<ref>Leven, Rachel. “U.S. Desalination Industry Grows Since 2000; See as Essential to Meeting Supply Needs”. Bloomberg BNA. August 21, 2013.</ref>

The profitability of the project is naturally heavily dependent on the price of water in California. If the drought ends, water supply will increase and the water price will drop significantly, but if the drought continues, the price of water in California could rise even more rapidly and allow desalination plants to be more profitable.<ref>Southern California Public Radio. “Where is California water use decreasing?”. December 2015.</ref> It may be difficult to make this plant profitable without a guaranteed purchase or other subsidy from the local governments because the presence of pre-existing plants maintains the water price relatively low. However, external sources of profit such as guaranteed purchase or subsidy are expected because the ramifications of not having desalination plant or other sources of water may be significantly more expensive than the subsidizing a desalination plant assuming the drought continues.<ref>Weiser, Matt. “Could desalination solve California’s water problem?”. The Sacramento Bee. October 18, 2014.</ref>

=Design Basis and Technical Approach=

[[File:Figure_2.jpg|left|200px|frame|Figure 2. Vertical Well Feed]]
Considering the number of pre-existing desalination plant and the potential risk of the drought ending, our plant will be tentatively designed to produce 15 million gallons of drinkable water per day to homes and businesses in the Los Angeles county area. With a recovery rate of near 30%, and a requirement for diluting any brine, the input feed flowrate will be 70.5 million gallons per day. Following the similarly built Carlsbad plant, we expect the plant to last 30 years. The project building timeline is one year based on our relatively small plant size. We will run for 350 days a year while shutting down for cleaning and maintenance once every 3 months (based on DOW standards).<ref name = "Bates">Bates, Wayne. “Cleaning your RO”. Hydranautics. </ref> The plant will be built 500 meters from the coastline on Long Beach. A borehole will be drilled 10 meters into the ground. At this depth there should be seawater aquifers below the coast that regularly recharges itself with seawater from the ocean. A vertical well containing intake holes will be put into the borehole. The intake holes will be equipped with large size filters that can easily be replaced. At the surface, the well will be connected to a large-scale pump in a well pump house. From there, the seawater will be horizontally pumped into a seawater tank in the desalination plant. This will be considered the start of our process. The process ends after pretreatment, desalination, post treatment, and brine treatment are completed. Appendix 1 shows more details on the Design Basis.

The feed composition is assumed to be the standard seawater “average” based on the DOW Chemical ROSA 9.0 Technical Manual in Table 1. However, in reality the seawater composition at the plant site will have to be chemically tested. The final product will be based on L.A. standards for drinking water quality (Table 1). The specifications in the table are upper limits that we cannot exceed. We will aim to reach the salt level standards but any remineralization and pH treatment will be exported and paid for due to the added complexity and cost in making such a system in-house.

Due to the acidity of the desalinated water coming out of the reverse osmosis (RO) system (pH ~ 6), the pumps and pipes carrying the desalinated water for post treatment will likely corrode over time. We have decided to place the desalination plant by the Long Beach Water Reclamation Plant (about 50 meters away) to prevent pipe corrosion.

[[File:Table_1.jpg|center|200px|frame|Table 1. Concentration of the standard seawater, the output from the RO system, and LA standard.]]

=Desalination Overview and Alternatives=

We first began our research into the current landscape of the desalination industry, particularly of plants that were built in the United States. We discovered there were many processes but the main ones included: 
* Membrane Separation (Forward Osmosis and Reverse Osmosis)
* Thermal Separation (Multi Stage Flash Distillation)
* Phase Change Separation (Freeze Separation)
Each process was analyzed for their pros and cons. For our project design, we decided to go with reverse osmosis as our desalination process. This is because it is the most commonly used process in the industry, and for good reason. Unlike thermal separation, which requires huge amounts of energy to supply heat, reverse osmosis is very energy efficient. This reduces energy costs that would be spread across the plant’s lifetime.<ref name = "Republic">Republic of Mauritius NY. "Technical Aspects of Desalination Plant". Sustainable Sanitation and Water Management. </ref>

Forward osmosis is a newer technology being developed that also uses membranes to separate water and salt and may be more energy efficient overall than reverse osmosis. This is because forward osmosis doesn’t require a pump to pressurize seawater through the membranes, but rather leaves a higher salt concentrated solution (AKA draw solution) on the other side of the membrane for the seawater to be drawn into. Then a second step is required to regenerate the concentrated draw solution and produce purified water.<ref>Mallinson, Alissa. "Study Shows Forward Osmosis Desalination Not Energy Efficient." MIT News [Boston]. July 23, 2014.</ref> We decided to move away from forward osmosis because it is a process that is still being refined and developed for desalination and lacks the data, programs, and technologies that reverse osmosis users benefit from. Additionally, a recent study from MIT has claimed that due to the inherent nature of forward osmosis requiring that seawater be drawn to a higher concentrated salt solution before the water and salt is separated, it cannot possibly be more energy efficient than reverse osmosis.<ref>O'Hern, Sean C., Karnik, Rohit. Nanofiltration across Defect-Sealed Nanoporous Monolayer Graphene. American Chemical Society 15.5 (2014); 3254-260. </ref>

Other processes such as freeze separation were too restrictive in their feed and process demands and required a lot of capital. Ultimately, reverse osmosis is simply the best process that exists for desalination and has been ranked the best when performance is evaluated in terms of technical, environmental, and economic aspects. It is a process that is flexible in water quantity, operation start-up and shut-off, and site location.

Nevertheless there are disadvantages to reverse osmosis that must be considered. The feed seawater must not be too salty and too low quality. The pressure to pump the seawater is enormous. The construction time is long for large plants. The capital costs are also high when compared with other processes such as thermal separation. But all of these disadvantages are ultimately outweighed by the advantages reverse osmosis provides.<ref name = "Republic"/>

=Process Overview and Alternatives=

See Appendix 2 (Figure 1) for the Process Flow Diagram (PFD). The process itself is detailed below. The stream flow rates are in Appendix 2 (Table 1). We assumed 10 meters of pipe for every stream except the inlet and outlet streams of the entire PFD.

==Pretreatment==

After the 10 meter vertical well pumps seawater from the Long Beach coast, the feed is pumped 500 meters to a seawater tank that makes it easy to shut-down the system and reduce the large volume of flow. The level control keeps the seawater from overfilling the tank. From there, part of the seawater goes to brine treatment while the rest go to pretreatment. The purpose of pretreatment in the desalination process is to prevent equipment damage by removing particles and microorganisms from the water. A variety of pollutants -- including calcium carbonate, sulfates, silica, clay, bacteria, and other colloids and microorganisms -- can disrupt the reverse osmosis process. Media filtration uses fine grains such as sand or anthracite to filter the particles. We used a large filter to catch larger particle material (e.g. sand). However, simple filtration is not enough to prevent colloid fouling. An ultrafilter (i.e. microfiltration) is used to remove virtually all suspended matter, effectively preventing colloid fouling. The costs associated with ultrafiltration are not prohibitive in terms of the project’s budget.<ref name = "Desalination">Desalination Pre-treatment. Lenntech.</ref>

Reverse osmosis systems can also be disrupted by the growth of microbiological films of bacteria on the film and pipe walls. They are typically treated as colloids during pretreatment, and most organisms are removed from the water by ultrafiltration, but this does not remove the smallest of organic particles. DBNPA is the most effective biocide, but these aren’t often used and have to be removed from the water. Chlorination is an effective way to remove bacteria, but it can cause degradation of the membrane, and must be removed before the reverse osmosis occurs. In our process we will use UV irradiation to remove any organic material that passes through the microfilters.<ref name = "Desalination"/>

Scaling can reduce the flux across the membrane and can build up in pipe corners, causing backup. It is caused by inorganic salts -- specifically sulfates and calcium and magnesium carbonate -- present in the water. Scaling can be prevented and/or controlled with acids, antiscalants, or ion exchange softening. Antiscalants are good for removing sulfates, but not much else. Acidification will only be effective at removing CaCO3 from water but, the decreased pH may deteriorate the membrane and pipes, so it may require future research.<ref name = "Desalination"/> Ion exchange is very expensive and requires more capital to undergo, but is overall the more effective option.

The ion exchanger is a pretreatment system which acts as a method to fight membrane scaling in reverse osmosis. The ion exchanger works by using sodium carboxylate as a resin which will react with Chlorine and Magnesium ions, removing them from the feed. The sodium is substituted out and enters the feed stream, increasing the overall sodium level and TDS. The lack of magnesium and chlorine ions means that there is nothing to react with the carbonates to form scaling salts on the membranes.

==Desalination (Reverse Osmosis)==

Two pumps (P-105 & P-106) take the adjusted feed seawater up to 58 bar. However, before this pumping occurs, about 70% of the original stream gets separated and heads into a PX Pressure Exchanger. This separated stream later joins together with the pumped stream to enter the reverse osmosis system at the required pressure of 58 bar. The streams can combine together as they have equal pressure. It is the concentrate stream from the output of the reverse osmosis system that transfers its energy to the adjusted feed in the PX Pressure Exchanger.

[[File:Figure_3.jpg|left|300px|frame|Figure 3. Inside look of pressure exchanger]]
The pressure of the concentrate stream is about 57 bar and it is fed into an exchanger which works by using vessel that resembles a small pipe within a rotating wheel. The low pressure (LP) feed enters the vessel at 1 bar (see Figure 3). The feed turns and is pushed out into an outlet pipe by the high pressure (HP) concentrate stream at ~57 bar. The feed is now a HP stream via this exchanged energy. The now LP concentrate stream turns and is pushed out into an outlet pipe by the LP feed entering the vessel.

98% of the energy within the concentrate stream can be recovered via the PX Pressure Exchanger.<ref>Desalination Product and Applications Catalog 2013. “Advance Technologies for Seawater Reverse Osmosis and Brackish Water Applications”. Energy Recovery, Inc. 2012-2014.</ref >Thus, the now HP feed stream is not quite 58 bar and will need to be pumped by a small booster pump (P-107).

When designing a reverse osmosis system, we need to make decisions on what specific kind of reverse osmosis system will be implemented. When making the decision, we investigated the advantages of different types of reverse osmosis among the following options: batch vs. continuous process, number of stages, and number of elements.<ref>FILMTEC Reverse Osmosis Membranes. “Water & Process Solutions Technical Manual.” Dow Chemical Company.</ref> Our decision was made based on process optimization which is described later in this report. The end result of that optimization was to go with a single stage, continuous process with five elements in series.

=Post Treatment=

The permeate is sent to post treatment after the salt content is analyzed and controlled if too high or too low due to any unexpected errors. As stated above, the post treatment process will be done 50 meters from our plant by the Long Beach Water Reclamation Plant. The reason we aren’t performing post treatment “in-house” is due to the increased cost it would take for us to build and run the equipment for post treatment. We have the option to use the government’s water treatment facilities that already exist and our 15 million gallons per day output is a relatively small addition to the large Long Beach Water Reclamation Plant. We will assume any costs to post treatment will be pushed to the consumer in the price of purchasing the drinking water from the Long Beach Water Reclamation Plant.

Post treatment at any facility works to bring water to “taste standards”. The water that is the permeate of our process is already meeting L.A. drinking standards as shown above in Table 1 after our mass balance calculations around the reverse osmosis system (Appendix 3, Table 1). The post treatment increases the pH from 6.3 to 7.3 with the addition of base and then adds minerals like fluoride, calcium, and magnesium back in to improve taste.

==Brine Treatment and Byproduct Waste==

A natural consequence of a reverse osmosis procedure is that a lot of salt will be leftover from the membrane separation in a very concentrated form. This is called “brine” and will have to be taken care of. According to EPA requirements, the salt concentration of water entering the environment (via land or sea) cannot be higher than 40,000 ppm.8 The salt concentration of our brine is much higher (around 70,000 ppm). Thus, the brine must be removed or treated. We decided to treat the brine by diluting it as removing brine can be difficult and may require lots of land (e.g. evaporation ponds).

After the concentrate exchanges energy in the energy exchanger, it becomes a low pressure stream. It is pumped into a dilution tank and mixes with 18.2 million gallons per day of seawater from the inlet to become diluted brine which is pumped 600 meters out to sea.

The waste byproducts are the solid buildup on the filters and membranes. The filters and membranes also have a set lifetime and thus will also need to be removed. The waste is all non-hazardous and thus can be put in the municipal dump.

=Process Details and Equipment Sizing=

==Filters and UV Treatment [Pretreatment]==
The cost of our ultrafiltration system is estimated using data from the University of Tennessee at Knoxville.<ref>Qiang He and R. Bruce Robinson Membrane Filtration Systems. “Membrane Module Costs.” University of Tennessee at Knoxville. December 2014.</ref> It estimated the cost of the cost of the ultrafiltration $0.057 per 1000 gallons and the construction and capital costs are $0.39 per gallon per day. The cost of UV treatment is based off of units available from buyultraviolet.com for units that can treat about 450 gallons per minute.<ref>Atlantic Ultraviolet Corporation. “Megatron Water Disinfection System.”</ref> Due to the water having to be in relatively small pipes for treatment we do not expect the costs to scale very much, so we believe these units to be an accurate cost estimate.

==Ion Exchanger [Pretreatment] ==

The ion exchanger was sized using data from DOW. Our process will use 60 vessels with ion exchange beds of volume 5.73 m3. With a flow rate of 30 bed volumes/hour (10,312.2 m3/hr) the system of beds has a capacity of 125% of the feed. This allows for periodic regeneration of the resin, a weak acid cation resin (DOW MAC-3 Resin). This weak cation resin will remove cationic calcium and magnesium from the water to prevent scaling. The mass flow was calculated via a mass balance on the ion exchanger (Appendix 3).

==Reverse Osmosis System [Desalination]==

Each of the 2607 Reverse Osmosis SW30HRLE-440i [Seawater 30 High Rejection Low Energy 440 ft2 interlocking cap membrane] units have the same dimensions. A, the net length, is 1016 mm. B, the total length including the endcap, is 1029 mm. C, the diameter of the filter element, is 201mm. D, the diameter of the permeate pathway, is 29 mm. For a single stage, we have five of the elements in series so the overall length of a single stage is expected to be roughly 5145 mm. The membrane material is polyamide thin film composite. It will last three years.
Operating Conditions [ROSA 9.0]: 
* Temperature: 30oC 
* Pressure: 58 bar 
* Flow Rate: 52.3 Million gal per day 
* Recovery: 28.7%

[[File:Figure_4.png|center|400px|frame|Figure 4. Diagram of Dow FILMTEC SW30HRLE-440i]]
Cleaning Schedule: 
* Conservative Estimate: Once every three months (three days available for cleaning)<ref name = "Bates"/> 
* DOW FILMTEC Cleaning Conditions 
** The normalized permeate flow drops 10%
* The normalized salt passage increases 5 - 10%

==Seawater Holding Tank and Dilution Tank [Brine Treatment]==

The initial seawater holding tank works as a buffer to prevent surges and help perform shutdowns quickly and easily. We expect to run continuously for 350 days a year with the rest of the available time spread out and used for shutting down the plant for cleaning membranes, filters, and the fixing of unexpected errors. There is a big tradeoff between having a holding tank with high retention time and having a holding tank with low cost. With that in mind, we erred on the side of lowering capital cost and went with a retention time of one hour. The volume could be found by multiplying the retention time with the feed flow rate (Stream 1). We assumed that a good height to width ratio for a large water tank would be 4:1.

A mass balance was done around the dilution tank to find what amount of seawater would be necessary to dilute the brine to 45,000 ppm salinity (Appendix 3; Table 3). With this the feed flow rate into the tank could be found. It was found that it takes about 20 minutes of retention time for good dilution of the brine in seawater.8 We again assumed that a good height to width ratio for a large water tank would be 4:1.

We assumed a corrosion allowance of 4 mm wall thickness. We calculated wall thickness using vessel pressure, hoop stress, longitudinal stress, and weld factor. Each tank is made with SS316 to protect against corrosion.<ref>Towler G, Sinnott R. Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013.</ref>

<center>
'''Table 2.''' Tank sizes.
[[File:Table_2.jpg|center|600px]]
</center>

=Process Optimization=
<center> '''Table 3.'''Optimization Criteria
[[File:Table_3.jpg|center|400px]]
</center>
In order to optimize the reverse osmosis system, numerous simulations were run. Three representative cases are reported as shown above (Table 3). The objective of the optimization was to minimize the required water price to break even in 10 years. In order to achieve this goal, the balance between the capital cost from the number of the filter elements and utility cost from energy consumption was weighed. Before the energy recovery system was implemented, the general trend was that if we have more filter elements in series, overall recovery increases but recovery rate per filter decreases, and we need more filter elements. While increasing the elements in series increase the required number of filters, and thereby increases the capital cost, we require less energy to produce the same amount of water. After the energy recovery system is implemented, the correlation between the number of elements in series and the required energy becomes more complicated because high pressure concentrate is used to pressurize the feed. However, the trade-off between the utility cost and capital cost still exists. Of the three cases shown above, single stage, five elements RO system turned out to be the most optimal solution because single stage, single element system required higher utility cost and two stage system slightly did not meet the LA standard. Detailed information which was used for the optimization is available in Appendix 3 and in Appendix 6.

=Ethical, Environmental, and Societal Considerations=

There are a few key considerations we have followed to meet ethical, environmental, and societal guidelines and requirements: 
* Meet L.A. Drinking Water Standards
** Correct ion concentration
* Moved near water treatment plant
** Pipes will not corrode from low pH
* Working below pressure limits
* Water will be locally served and meet taste standards
** Via Post Treatment mineralization
* Follows EPA Environmental Salinity Requirements
* Reduced Energy Requirements
** PX Pressure Exchanger
** Optimized RO System
** Reverse Osmosis is itself Energy Saving

For further environmental considerations, see effluent summary in Appendix 4.

=Economic Analysis=

The pricing of the majority of the process was completed using the Aspen Economic Analyzer, however we had determine costs for the ultrafiltration, reverse osmosis membranes, and the ultraviolet treatment was completed separately. This led to a total ISBL of $56 MM, OSBL of $22.4 MM, Engineering Costs of $22.4 MM and a Contingency of $7.8 MM. We also chose a working Capital of $10 MM. This working capital is sufficient because we have no raw materials costs. We also had FCOP of $5.2 MM and VCOP of $3.5 MM.

We originally were planning on pricing our water at $700 per acre-foot, but after doing economic analysis we had a 10 year NPV of $-76.2 MM. After determining did a sensitivity analysis and determined that if our water was priced at $2104 per acre-foot we could have an NPV of $0 after 10 years and if we were able to get the same price of $2257 per acre-foot as the plant in Carlsbad we would have a 10 year NPV of $8.3 MM which increases to $41.1 MM at 20 years. For more detailed data about the economic analysis see Appendix 6.

=Conclusion and Recommendations=

Our mission was to create an economically viable process that will desalinate water for human consumption in the L.A. County. We have found after research and optimization that the best process to undergo desalination is a continuous single stage, five elements reverse osmosis system. Our process can meet the required L.A. drinking water standard and we have set up plans for post treatment at a local water reclamation site. While meeting the drinking water standards, we also met our ethical, environmental, and social standards.

Nevertheless, at a sale price of $700 per acre-foot, our proposed desalination plant in Los Angeles, California looks to be a very expensive project to undergo. Without any sort of government assistance to carry the heavy capital costs of such a project, we would suggest the plant not be built. However, after performing a sensitivity analysis at higher (and less conservative) price points we reached a breakeven point in 10 years with a sale price of $2104 per acre-foot of drinking water. Research has shown that the droughts striking California look only to get worse in the near future. With limited freshwater resources, there is no option other than to consider desalination. Therefore it is reasonable to expect we could get near the same price deal as local desalination plants like the Carlsbad Plant ($2257 per acre-foot) As we have met all design constraints for this project, we hope to convince policymakers to adopt our project at least $2104 per acre-foot water. If this deal is reached, we recommend that this project and design be funded.

However, there are ways to improve our design with future work. Right now all the pipes diameters are based on having a flowrate of 1.5 m/s. But in reality, we’d need to use standard piping sizes to get more accurate pricing and energy costs. Also, in reality we would be purchasing some of these vessels in bulk quantities (i.e. bulk filters) so the price could be lowered with negotiation and economies of scale. We also have pumps that are providing very little pressure for their high cost. We can remove these pumps and simply run at higher pressures earlier. We can potentially reduce early energy costs by building a “floating” desalination plant in the sea to reduce pumping distance from the well.

=Appendices=

==Appendix 1: Design Basis==
[[File:Appendix_1.png|thumb|border|center|800px|middle|]]

==Appendix 2: Process Flow Diagram==
[[File:Appendix_2.png|center]]
==Appendix 3: Mass and Energy Balance Tables==
'''Table 1'''. Mass balance on 1 stage 5 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T1.png|center|600px|]]

'''Table 2'''. Mass balance on 1 stage 1 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T2.png|center|600px|]]

'''Table 3'''. Mass balance on 2 stage 5,5 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T3.png|center|600px|]]

'''Table 4'''. 1 stage 5 elements CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T4.png|center|400px|]]

'''Table 5'''. 1 stage 1 element CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T5.png|center|400px|]]

'''Table 6'''. 2 stages 5,5 elements CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T6.png|center|400px|]]

'''Table 7'''. Energy balance on 1 stage 5 elements RO system
[[File:Appendix_3_T7.png|center|400px|]]

==Appendix 4: Effluent Summary==
[[File:Appendix_4.png|center]]

==Appendix 5: Cost of Production==
[[File:Appendix_5.png|center]]

==Appendix 6: Economic Anlaysis==
===Case I ===
Water is $700; RO with 1 stage and 5 elements
[[File:Appendix_6_C1.png|center]]

===Case II===
Water is $1400; RO with 1 stage and 5 elements
[[File:Appendix_6_C2.png|center]]

===Case III: Carlsbad ===
Water is $2257; RO with 1 stage and 5 elements
[[File:Appendix_6_C3.png|center]]

===Case IV: Breakeven (NPV = 0 at 10 years)===
Water is $2104; RO with 1 stage and 5 elements
[[File:Appendix_6_C4.png|center]]

===Case V: Breakeven (NPV = 0 at 10 years)===
Water is $2112; RO with 2 stages and 5,5 elements
[[File:Appendix_6_C5.png|center]]

===Case VI: Breakeven (NPV = 0 at 10 years)===
Water is $2138 with 1 stage and 1 element
[[File:Appendix_6_C6.png|center]]

==Appendix 7: ROSA Simulation Outputs==

'''Table 1'''. System detail of 1 stage 5 elements RO system
[[File:Appendix_7_T1.png|center]]

'''Table 2'''. System detail of 1 stage 1 elements RO system
[[File:Appendix_7_T2.png|center]]

'''Table 3'''. System detail of 2 stage 5,5 elements system
[[File:Appendix_7_T3.png|center]]

'''Table 4'''. 1 stage 5 elements RO system concentrations and pH in each stream
[[File:Appendix_7_T4.png|center]]

'''Table 5'''. 1 stage 1 element RO system concentration and pH in each stream
[[File:Appendix_7_T5.png|center]]

'''Table 6'''. 2 stage 5,5 elements RO system concentrations and pH in each stream
[[File:Appendix_7_T6.png|center]]

=References=
<references/>

Desalination - Team E

2016-03-12T03:32:21Z

Bts511: /* Appendix 3: Mass and Energy Balance Tables */

Group E Corporation

Authors: Hassan Ali, Woo Soo Choe, Brett Sleyster, Jake Stolley

Instructors: Fengqi You, David Wegerer

March 11, 2016

=Executive Summary=

Outlined in this report is the project model report for the proposed desalination plant in Los Angeles, California which will provide drinking water to local citizens using reverse osmosis process. This location was chosen due to the region’s propensity for droughts in recent years. The fresh water resources in California have plummeted as water prices have risen 3-4% each year. This is only expected to continue. Other companies have worked deals with the Californian government to create desalination plants that provide a new fresh water drinking source. This suggests that there is a strong market here with potential growth.

The plant will have an output of 15 million gallons per day, which is about medium size relative to other plants. Thus the build time will be about one year. The plant will run for 30 years, with 350 days per year of use. Shutdown time will be used for cleaning and maintenance of equipment. The plant will be built on Long Beach, 500 meters from the coast nearby the Long Beach Water Reclamation Plant (LBWP). The input will withdraw 70.5 million gallons per day of seawater that will be used for reverse osmosis and dilution of brine.

Reverse osmosis was chosen as the desalination method for its low energy cost and flexibility. We designed a process that will pretreat the seawater to remove waste, colloids, bacteria, and membrane-damaging scalants. Once this seawater is treated is pumped to heavy pressures through a reverse osmosis system. After doing process optimization under various system parameter changes, we found a continuous, single stage, five element reverse osmosis process to be best. It produces a near 30% recovery. We reduce energy costs by including a pressure exchanger between our concentrate outlet and adjusted feed into the reverse osmosis system.

After performing mass and energy balances, we were able to size equipment and find our permeate composition. It turns out our process will meet L.A. standards for drinking water quality. However, post treatment will be done by the local LBWP to make the water taste better and be non-caustic to pipes. This will reduce “in-house” costs. Many environmental considerations were made, to meet our environmental standards. We meet EPA salinity requirements for our brine waste by diluting it. We also don’t use a lot of energy and thus reduce our fossil fuel consumption and CO2 generation. To be safe, we work well under max pressure limits for our equipment. We make sure we are ethical and meet all standards that our local county requires of us.

After an economic evaluation at a conservative well-water price of $700 per acre-foot, we realize that our revenue is way too low to cover our costs. Capital costs are especially large with regards to our large water tanks. We find that we reach a 10 year NPV breakpoint of $2104 per acre-foot after performing a sensitivity analysis. The Carlsbad Plant in San Diego worked out a deal for $2257 per acre-foot.

As we have met all design constraints for this project, we hope to convince policymakers to adopt our project at least $2104 per acre-foot water. If this deal is reached, we recommend that this project and design be funded. There is future work to be done to improve costs such as performing pump optimization. However, this work is minimal and won’t harm our bottom line if we reach our proposed price points.

__TOC__

=Introduction=

Due to a combination of various factors--including rising human population, climate change, increased energy consumption, and land erosion--the availability of fresh water as a resource for human consumption has been dwindling. This has had a negative impact on the world, but this impact has been especially felt in areas with natural shortages in fresh water. Desert regions, such as Saudi Arabia, have resorted to desalinating seawater at the coast and transporting up to their cities. This encompasses their entire freshwater system as very few natural water sources remain. In the United States there is a region which has shown increasing signs of following in the footsteps of Saudi Arabia and becoming fully reliant on desalination: Los Angeles, California.

=Location and Market Analysis=

In this project, Los Angeles was chosen as the location our plant, because the area has been subject to a severe drought for the last four years, as shown in Figure 1. The rapid decrease in the available water has caused permanent damages to the ground, decreased the amount of water in the reservoir from 45% to 25%, and reduced the amount of hydroelectric power produced in the state. Such deficiency in usable water has increased the water prices in Los Angeles by 6-7% annually for the last decade and is expected to rise by 3.4% annually for typical customers for the next five years. For customers who use more than 20000 gallons a year, the bill is expected to increase by 34% by 2021. This price increase is occurring because of both drought and maintenance. <ref>Pacific Institute: The California Drought. 2014.</ref>

[[File:Figure_1.jpg|right|200px|frame|Figure 1. California Drought Index.]]
As of March 2015 the Metropolitan Water District of Southern California was willing to pay as much as $700 per acre-foot of water to farmers in the Sacramento Valley, so we assume this will be the target price of the water produced by the this desalination plant.<ref>Hessel, Phil. “Dry Southern California Offers Northern Farmers Top Dollar for Water”. NBC News. March 18, 2015. </ref>

Currently, there is a plant in Carlsbad which supplies drinking water to the San Diego area. The plant can produce up to 50 million gallons per day and is the nation's largest water desalination plant. The water from this plant costs $2257 per acre-foot for the first 48000 acre-feet and about $2000 for any additional acre-foot. The county signed a contract with this plant to purchase the first 48000 acre-feet of water for the next 30 years.<ref>Carlsbad Desalination Plant. “Project Overview”. January 2016.</ref> Also, there is another desalination plant in Tampa that produces about 25 million gallons of water per day, and over 2000 other public and industrial plants with production capacity of over 300000 gallons per day. When conducting the economic analysis our plant, such numbers presented by other plants must be considered, so our project can yield positive profit.<ref>Leven, Rachel. “U.S. Desalination Industry Grows Since 2000; See as Essential to Meeting Supply Needs”. Bloomberg BNA. August 21, 2013.</ref>

The profitability of the project is naturally heavily dependent on the price of water in California. If the drought ends, water supply will increase and the water price will drop significantly, but if the drought continues, the price of water in California could rise even more rapidly and allow desalination plants to be more profitable.<ref>Southern California Public Radio. “Where is California water use decreasing?”. December 2015.</ref> It may be difficult to make this plant profitable without a guaranteed purchase or other subsidy from the local governments because the presence of pre-existing plants maintains the water price relatively low. However, external sources of profit such as guaranteed purchase or subsidy are expected because the ramifications of not having desalination plant or other sources of water may be significantly more expensive than the subsidizing a desalination plant assuming the drought continues.<ref>Weiser, Matt. “Could desalination solve California’s water problem?”. The Sacramento Bee. October 18, 2014.</ref>

=Design Basis and Technical Approach=

[[File:Figure_2.jpg|left|200px|frame|Figure 2. Vertical Well Feed]]
Considering the number of pre-existing desalination plant and the potential risk of the drought ending, our plant will be tentatively designed to produce 15 million gallons of drinkable water per day to homes and businesses in the Los Angeles county area. With a recovery rate of near 30%, and a requirement for diluting any brine, the input feed flowrate will be 70.5 million gallons per day. Following the similarly built Carlsbad plant, we expect the plant to last 30 years. The project building timeline is one year based on our relatively small plant size. We will run for 350 days a year while shutting down for cleaning and maintenance once every 3 months (based on DOW standards).<ref name = "Bates">Bates, Wayne. “Cleaning your RO”. Hydranautics. </ref> The plant will be built 500 meters from the coastline on Long Beach. A borehole will be drilled 10 meters into the ground. At this depth there should be seawater aquifers below the coast that regularly recharges itself with seawater from the ocean. A vertical well containing intake holes will be put into the borehole. The intake holes will be equipped with large size filters that can easily be replaced. At the surface, the well will be connected to a large-scale pump in a well pump house. From there, the seawater will be horizontally pumped into a seawater tank in the desalination plant. This will be considered the start of our process. The process ends after pretreatment, desalination, post treatment, and brine treatment are completed. Appendix 1 shows more details on the Design Basis.

The feed composition is assumed to be the standard seawater “average” based on the DOW Chemical ROSA 9.0 Technical Manual in Table 1. However, in reality the seawater composition at the plant site will have to be chemically tested. The final product will be based on L.A. standards for drinking water quality (Table 1). The specifications in the table are upper limits that we cannot exceed. We will aim to reach the salt level standards but any remineralization and pH treatment will be exported and paid for due to the added complexity and cost in making such a system in-house.

Due to the acidity of the desalinated water coming out of the reverse osmosis (RO) system (pH ~ 6), the pumps and pipes carrying the desalinated water for post treatment will likely corrode over time. We have decided to place the desalination plant by the Long Beach Water Reclamation Plant (about 50 meters away) to prevent pipe corrosion.

[[File:Table_1.jpg|center|200px|frame|Table 1. Concentration of the standard seawater, the output from the RO system, and LA standard.]]

=Desalination Overview and Alternatives=

We first began our research into the current landscape of the desalination industry, particularly of plants that were built in the United States. We discovered there were many processes but the main ones included: 
* Membrane Separation (Forward Osmosis and Reverse Osmosis)
* Thermal Separation (Multi Stage Flash Distillation)
* Phase Change Separation (Freeze Separation)
Each process was analyzed for their pros and cons. For our project design, we decided to go with reverse osmosis as our desalination process. This is because it is the most commonly used process in the industry, and for good reason. Unlike thermal separation, which requires huge amounts of energy to supply heat, reverse osmosis is very energy efficient. This reduces energy costs that would be spread across the plant’s lifetime.<ref name = "Republic">Republic of Mauritius NY. "Technical Aspects of Desalination Plant". Sustainable Sanitation and Water Management. </ref>

Forward osmosis is a newer technology being developed that also uses membranes to separate water and salt and may be more energy efficient overall than reverse osmosis. This is because forward osmosis doesn’t require a pump to pressurize seawater through the membranes, but rather leaves a higher salt concentrated solution (AKA draw solution) on the other side of the membrane for the seawater to be drawn into. Then a second step is required to regenerate the concentrated draw solution and produce purified water.<ref>Mallinson, Alissa. "Study Shows Forward Osmosis Desalination Not Energy Efficient." MIT News [Boston]. July 23, 2014.</ref> We decided to move away from forward osmosis because it is a process that is still being refined and developed for desalination and lacks the data, programs, and technologies that reverse osmosis users benefit from. Additionally, a recent study from MIT has claimed that due to the inherent nature of forward osmosis requiring that seawater be drawn to a higher concentrated salt solution before the water and salt is separated, it cannot possibly be more energy efficient than reverse osmosis.<ref>O'Hern, Sean C., Karnik, Rohit. Nanofiltration across Defect-Sealed Nanoporous Monolayer Graphene. American Chemical Society 15.5 (2014); 3254-260. </ref>

Other processes such as freeze separation were too restrictive in their feed and process demands and required a lot of capital. Ultimately, reverse osmosis is simply the best process that exists for desalination and has been ranked the best when performance is evaluated in terms of technical, environmental, and economic aspects. It is a process that is flexible in water quantity, operation start-up and shut-off, and site location.

Nevertheless there are disadvantages to reverse osmosis that must be considered. The feed seawater must not be too salty and too low quality. The pressure to pump the seawater is enormous. The construction time is long for large plants. The capital costs are also high when compared with other processes such as thermal separation. But all of these disadvantages are ultimately outweighed by the advantages reverse osmosis provides.<ref name = "Republic"/>

=Process Overview and Alternatives=

See Appendix 2 (Figure 1) for the Process Flow Diagram (PFD). The process itself is detailed below. The stream flow rates are in Appendix 2 (Table 1). We assumed 10 meters of pipe for every stream except the inlet and outlet streams of the entire PFD.

==Pretreatment==

After the 10 meter vertical well pumps seawater from the Long Beach coast, the feed is pumped 500 meters to a seawater tank that makes it easy to shut-down the system and reduce the large volume of flow. The level control keeps the seawater from overfilling the tank. From there, part of the seawater goes to brine treatment while the rest go to pretreatment. The purpose of pretreatment in the desalination process is to prevent equipment damage by removing particles and microorganisms from the water. A variety of pollutants -- including calcium carbonate, sulfates, silica, clay, bacteria, and other colloids and microorganisms -- can disrupt the reverse osmosis process. Media filtration uses fine grains such as sand or anthracite to filter the particles. We used a large filter to catch larger particle material (e.g. sand). However, simple filtration is not enough to prevent colloid fouling. An ultrafilter (i.e. microfiltration) is used to remove virtually all suspended matter, effectively preventing colloid fouling. The costs associated with ultrafiltration are not prohibitive in terms of the project’s budget.<ref name = "Desalination">Desalination Pre-treatment. Lenntech.</ref>

Reverse osmosis systems can also be disrupted by the growth of microbiological films of bacteria on the film and pipe walls. They are typically treated as colloids during pretreatment, and most organisms are removed from the water by ultrafiltration, but this does not remove the smallest of organic particles. DBNPA is the most effective biocide, but these aren’t often used and have to be removed from the water. Chlorination is an effective way to remove bacteria, but it can cause degradation of the membrane, and must be removed before the reverse osmosis occurs. In our process we will use UV irradiation to remove any organic material that passes through the microfilters.<ref name = "Desalination"/>

Scaling can reduce the flux across the membrane and can build up in pipe corners, causing backup. It is caused by inorganic salts -- specifically sulfates and calcium and magnesium carbonate -- present in the water. Scaling can be prevented and/or controlled with acids, antiscalants, or ion exchange softening. Antiscalants are good for removing sulfates, but not much else. Acidification will only be effective at removing CaCO3 from water but, the decreased pH may deteriorate the membrane and pipes, so it may require future research.<ref name = "Desalination"/> Ion exchange is very expensive and requires more capital to undergo, but is overall the more effective option.

The ion exchanger is a pretreatment system which acts as a method to fight membrane scaling in reverse osmosis. The ion exchanger works by using sodium carboxylate as a resin which will react with Chlorine and Magnesium ions, removing them from the feed. The sodium is substituted out and enters the feed stream, increasing the overall sodium level and TDS. The lack of magnesium and chlorine ions means that there is nothing to react with the carbonates to form scaling salts on the membranes.

==Desalination (Reverse Osmosis)==

Two pumps (P-105 & P-106) take the adjusted feed seawater up to 58 bar. However, before this pumping occurs, about 70% of the original stream gets separated and heads into a PX Pressure Exchanger. This separated stream later joins together with the pumped stream to enter the reverse osmosis system at the required pressure of 58 bar. The streams can combine together as they have equal pressure. It is the concentrate stream from the output of the reverse osmosis system that transfers its energy to the adjusted feed in the PX Pressure Exchanger.

[[File:Figure_3.jpg|left|300px|frame|Figure 3. Inside look of pressure exchanger]]
The pressure of the concentrate stream is about 57 bar and it is fed into an exchanger which works by using vessel that resembles a small pipe within a rotating wheel. The low pressure (LP) feed enters the vessel at 1 bar (see Figure 3). The feed turns and is pushed out into an outlet pipe by the high pressure (HP) concentrate stream at ~57 bar. The feed is now a HP stream via this exchanged energy. The now LP concentrate stream turns and is pushed out into an outlet pipe by the LP feed entering the vessel.

98% of the energy within the concentrate stream can be recovered via the PX Pressure Exchanger.<ref>Desalination Product and Applications Catalog 2013. “Advance Technologies for Seawater Reverse Osmosis and Brackish Water Applications”. Energy Recovery, Inc. 2012-2014.</ref >Thus, the now HP feed stream is not quite 58 bar and will need to be pumped by a small booster pump (P-107).

When designing a reverse osmosis system, we need to make decisions on what specific kind of reverse osmosis system will be implemented. When making the decision, we investigated the advantages of different types of reverse osmosis among the following options: batch vs. continuous process, number of stages, and number of elements.<ref>FILMTEC Reverse Osmosis Membranes. “Water & Process Solutions Technical Manual.” Dow Chemical Company.</ref> Our decision was made based on process optimization which is described later in this report. The end result of that optimization was to go with a single stage, continuous process with five elements in series.

=Post Treatment=

The permeate is sent to post treatment after the salt content is analyzed and controlled if too high or too low due to any unexpected errors. As stated above, the post treatment process will be done 50 meters from our plant by the Long Beach Water Reclamation Plant. The reason we aren’t performing post treatment “in-house” is due to the increased cost it would take for us to build and run the equipment for post treatment. We have the option to use the government’s water treatment facilities that already exist and our 15 million gallons per day output is a relatively small addition to the large Long Beach Water Reclamation Plant. We will assume any costs to post treatment will be pushed to the consumer in the price of purchasing the drinking water from the Long Beach Water Reclamation Plant.

Post treatment at any facility works to bring water to “taste standards”. The water that is the permeate of our process is already meeting L.A. drinking standards as shown above in Table 1 after our mass balance calculations around the reverse osmosis system (Appendix 3, Table 1). The post treatment increases the pH from 6.3 to 7.3 with the addition of base and then adds minerals like fluoride, calcium, and magnesium back in to improve taste.

==Brine Treatment and Byproduct Waste==

A natural consequence of a reverse osmosis procedure is that a lot of salt will be leftover from the membrane separation in a very concentrated form. This is called “brine” and will have to be taken care of. According to EPA requirements, the salt concentration of water entering the environment (via land or sea) cannot be higher than 40,000 ppm.8 The salt concentration of our brine is much higher (around 70,000 ppm). Thus, the brine must be removed or treated. We decided to treat the brine by diluting it as removing brine can be difficult and may require lots of land (e.g. evaporation ponds).

After the concentrate exchanges energy in the energy exchanger, it becomes a low pressure stream. It is pumped into a dilution tank and mixes with 18.2 million gallons per day of seawater from the inlet to become diluted brine which is pumped 600 meters out to sea.

The waste byproducts are the solid buildup on the filters and membranes. The filters and membranes also have a set lifetime and thus will also need to be removed. The waste is all non-hazardous and thus can be put in the municipal dump.

=Process Details and Equipment Sizing=

==Filters and UV Treatment [Pretreatment]==
The cost of our ultrafiltration system is estimated using data from the University of Tennessee at Knoxville.<ref>Qiang He and R. Bruce Robinson Membrane Filtration Systems. “Membrane Module Costs.” University of Tennessee at Knoxville. December 2014.</ref> It estimated the cost of the cost of the ultrafiltration $0.057 per 1000 gallons and the construction and capital costs are $0.39 per gallon per day. The cost of UV treatment is based off of units available from buyultraviolet.com for units that can treat about 450 gallons per minute.<ref>Atlantic Ultraviolet Corporation. “Megatron Water Disinfection System.”</ref> Due to the water having to be in relatively small pipes for treatment we do not expect the costs to scale very much, so we believe these units to be an accurate cost estimate.

==Ion Exchanger [Pretreatment] ==

The ion exchanger was sized using data from DOW. Our process will use 60 vessels with ion exchange beds of volume 5.73 m3. With a flow rate of 30 bed volumes/hour (10,312.2 m3/hr) the system of beds has a capacity of 125% of the feed. This allows for periodic regeneration of the resin, a weak acid cation resin (DOW MAC-3 Resin). This weak cation resin will remove cationic calcium and magnesium from the water to prevent scaling. The mass flow was calculated via a mass balance on the ion exchanger (Appendix 3).

==Reverse Osmosis System [Desalination]==

Each of the 2607 Reverse Osmosis SW30HRLE-440i [Seawater 30 High Rejection Low Energy 440 ft2 interlocking cap membrane] units have the same dimensions. A, the net length, is 1016 mm. B, the total length including the endcap, is 1029 mm. C, the diameter of the filter element, is 201mm. D, the diameter of the permeate pathway, is 29 mm. For a single stage, we have five of the elements in series so the overall length of a single stage is expected to be roughly 5145 mm. The membrane material is polyamide thin film composite. It will last three years.
Operating Conditions [ROSA 9.0]: 
* Temperature: 30oC 
* Pressure: 58 bar 
* Flow Rate: 52.3 Million gal per day 
* Recovery: 28.7%

[[File:Figure_4.png|center|400px|frame|Figure 4. Diagram of Dow FILMTEC SW30HRLE-440i]]
Cleaning Schedule: 
* Conservative Estimate: Once every three months (three days available for cleaning)<ref name = "Bates"/> 
* DOW FILMTEC Cleaning Conditions 
** The normalized permeate flow drops 10%
* The normalized salt passage increases 5 - 10%

==Seawater Holding Tank and Dilution Tank [Brine Treatment]==

The initial seawater holding tank works as a buffer to prevent surges and help perform shutdowns quickly and easily. We expect to run continuously for 350 days a year with the rest of the available time spread out and used for shutting down the plant for cleaning membranes, filters, and the fixing of unexpected errors. There is a big tradeoff between having a holding tank with high retention time and having a holding tank with low cost. With that in mind, we erred on the side of lowering capital cost and went with a retention time of one hour. The volume could be found by multiplying the retention time with the feed flow rate (Stream 1). We assumed that a good height to width ratio for a large water tank would be 4:1.

A mass balance was done around the dilution tank to find what amount of seawater would be necessary to dilute the brine to 45,000 ppm salinity (Appendix 3; Table 3). With this the feed flow rate into the tank could be found. It was found that it takes about 20 minutes of retention time for good dilution of the brine in seawater.8 We again assumed that a good height to width ratio for a large water tank would be 4:1.

We assumed a corrosion allowance of 4 mm wall thickness. We calculated wall thickness using vessel pressure, hoop stress, longitudinal stress, and weld factor. Each tank is made with SS316 to protect against corrosion.<ref>Towler G, Sinnott R. Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013.</ref>

<center>
'''Table 2.''' Tank sizes.
[[File:Table_2.jpg|center|600px]]
</center>

=Process Optimization=
<center> '''Table 3.'''Optimization Criteria
[[File:Table_3.jpg|center|400px]]
</center>
In order to optimize the reverse osmosis system, numerous simulations were run. Three representative cases are reported as shown above (Table 3). The objective of the optimization was to minimize the required water price to break even in 10 years. In order to achieve this goal, the balance between the capital cost from the number of the filter elements and utility cost from energy consumption was weighed. Before the energy recovery system was implemented, the general trend was that if we have more filter elements in series, overall recovery increases but recovery rate per filter decreases, and we need more filter elements. While increasing the elements in series increase the required number of filters, and thereby increases the capital cost, we require less energy to produce the same amount of water. After the energy recovery system is implemented, the correlation between the number of elements in series and the required energy becomes more complicated because high pressure concentrate is used to pressurize the feed. However, the trade-off between the utility cost and capital cost still exists. Of the three cases shown above, single stage, five elements RO system turned out to be the most optimal solution because single stage, single element system required higher utility cost and two stage system slightly did not meet the LA standard. Detailed information which was used for the optimization is available in Appendix 3 and in Appendix 6.

=Ethical, Environmental, and Societal Considerations=

There are a few key considerations we have followed to meet ethical, environmental, and societal guidelines and requirements: 
* Meet L.A. Drinking Water Standards
** Correct ion concentration
* Moved near water treatment plant
** Pipes will not corrode from low pH
* Working below pressure limits
* Water will be locally served and meet taste standards
** Via Post Treatment mineralization
* Follows EPA Environmental Salinity Requirements
* Reduced Energy Requirements
** PX Pressure Exchanger
** Optimized RO System
** Reverse Osmosis is itself Energy Saving

For further environmental considerations, see effluent summary in Appendix 4.

=Economic Analysis=

The pricing of the majority of the process was completed using the Aspen Economic Analyzer, however we had determine costs for the ultrafiltration, reverse osmosis membranes, and the ultraviolet treatment was completed separately. This led to a total ISBL of $56 MM, OSBL of $22.4 MM, Engineering Costs of $22.4 MM and a Contingency of $7.8 MM. We also chose a working Capital of $10 MM. This working capital is sufficient because we have no raw materials costs. We also had FCOP of $5.2 MM and VCOP of $3.5 MM.

We originally were planning on pricing our water at $700 per acre-foot, but after doing economic analysis we had a 10 year NPV of $-76.2 MM. After determining did a sensitivity analysis and determined that if our water was priced at $2104 per acre-foot we could have an NPV of $0 after 10 years and if we were able to get the same price of $2257 per acre-foot as the plant in Carlsbad we would have a 10 year NPV of $8.3 MM which increases to $41.1 MM at 20 years. For more detailed data about the economic analysis see Appendix 6.

=Conclusion and Recommendations=

Our mission was to create an economically viable process that will desalinate water for human consumption in the L.A. County. We have found after research and optimization that the best process to undergo desalination is a continuous single stage, five elements reverse osmosis system. Our process can meet the required L.A. drinking water standard and we have set up plans for post treatment at a local water reclamation site. While meeting the drinking water standards, we also met our ethical, environmental, and social standards.

Nevertheless, at a sale price of $700 per acre-foot, our proposed desalination plant in Los Angeles, California looks to be a very expensive project to undergo. Without any sort of government assistance to carry the heavy capital costs of such a project, we would suggest the plant not be built. However, after performing a sensitivity analysis at higher (and less conservative) price points we reached a breakeven point in 10 years with a sale price of $2104 per acre-foot of drinking water. Research has shown that the droughts striking California look only to get worse in the near future. With limited freshwater resources, there is no option other than to consider desalination. Therefore it is reasonable to expect we could get near the same price deal as local desalination plants like the Carlsbad Plant ($2257 per acre-foot) As we have met all design constraints for this project, we hope to convince policymakers to adopt our project at least $2104 per acre-foot water. If this deal is reached, we recommend that this project and design be funded.

However, there are ways to improve our design with future work. Right now all the pipes diameters are based on having a flowrate of 1.5 m/s. But in reality, we’d need to use standard piping sizes to get more accurate pricing and energy costs. Also, in reality we would be purchasing some of these vessels in bulk quantities (i.e. bulk filters) so the price could be lowered with negotiation and economies of scale. We also have pumps that are providing very little pressure for their high cost. We can remove these pumps and simply run at higher pressures earlier. We can potentially reduce early energy costs by building a “floating” desalination plant in the sea to reduce pumping distance from the well.

=Appendices=

==Appendix 1: Design Basis==
[[File:Appendix_1.png|thumb|border|center|800px|middle|]]

==Appendix 2: Process Flow Diagram==
[[File:Appendix_2.png|center]]
==Appendix 3: Mass and Energy Balance Tables==
'''Table 1'''. Mass balance on 1 stage 5 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T1.png|center|600px|]]

'''Table 2'''. Mass balance on 1 stage 1 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T2.png|center|600px|]]

'''Table 3'''. Mass balance on 2 stage 5,5 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T3.png|center|600px|]]

'''Table 4'''. 1 stage 5 elements CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T4.png|center|300px|]]

'''Table 5'''. 1 stage 1 element CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T5.png|center|300px|]]

'''Table 6'''. 2 stages 5,5 elements CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T6.png|center|300px|]]

'''Table 7'''. Energy balance on 1 stage 5 elements RO system
[[File:Appendix_3_T7.png|center|300px|]]

==Appendix 4: Effluent Summary==
[[File:Appendix_4.png|center]]

==Appendix 5: Cost of Production==
[[File:Appendix_5.png|center]]

==Appendix 6: Economic Anlaysis==
===Case I ===
Water is $700; RO with 1 stage and 5 elements
[[File:Appendix_6_C1.png|center]]

===Case II===
Water is $1400; RO with 1 stage and 5 elements
[[File:Appendix_6_C2.png|center]]

===Case III: Carlsbad ===
Water is $2257; RO with 1 stage and 5 elements
[[File:Appendix_6_C3.png|center]]

===Case IV: Breakeven (NPV = 0 at 10 years)===
Water is $2104; RO with 1 stage and 5 elements
[[File:Appendix_6_C4.png|center]]

===Case V: Breakeven (NPV = 0 at 10 years)===
Water is $2112; RO with 2 stages and 5,5 elements
[[File:Appendix_6_C5.png|center]]

===Case VI: Breakeven (NPV = 0 at 10 years)===
Water is $2138 with 1 stage and 1 element
[[File:Appendix_6_C6.png|center]]

==Appendix 7: ROSA Simulation Outputs==

'''Table 1'''. System detail of 1 stage 5 elements RO system
[[File:Appendix_7_T1.png|center]]

'''Table 2'''. System detail of 1 stage 1 elements RO system
[[File:Appendix_7_T2.png|center]]

'''Table 3'''. System detail of 2 stage 5,5 elements system
[[File:Appendix_7_T3.png|center]]

'''Table 4'''. 1 stage 5 elements RO system concentrations and pH in each stream
[[File:Appendix_7_T4.png|center]]

'''Table 5'''. 1 stage 1 element RO system concentration and pH in each stream
[[File:Appendix_7_T5.png|center]]

'''Table 6'''. 2 stage 5,5 elements RO system concentrations and pH in each stream
[[File:Appendix_7_T6.png|center]]

=References=
<references/>

Desalination - Team E

2016-03-12T03:31:40Z

Bts511: /* Appendix 3: Mass and Energy Balance Tables */

Group E Corporation

Authors: Hassan Ali, Woo Soo Choe, Brett Sleyster, Jake Stolley

Instructors: Fengqi You, David Wegerer

March 11, 2016

=Executive Summary=

Outlined in this report is the project model report for the proposed desalination plant in Los Angeles, California which will provide drinking water to local citizens using reverse osmosis process. This location was chosen due to the region’s propensity for droughts in recent years. The fresh water resources in California have plummeted as water prices have risen 3-4% each year. This is only expected to continue. Other companies have worked deals with the Californian government to create desalination plants that provide a new fresh water drinking source. This suggests that there is a strong market here with potential growth.

The plant will have an output of 15 million gallons per day, which is about medium size relative to other plants. Thus the build time will be about one year. The plant will run for 30 years, with 350 days per year of use. Shutdown time will be used for cleaning and maintenance of equipment. The plant will be built on Long Beach, 500 meters from the coast nearby the Long Beach Water Reclamation Plant (LBWP). The input will withdraw 70.5 million gallons per day of seawater that will be used for reverse osmosis and dilution of brine.

Reverse osmosis was chosen as the desalination method for its low energy cost and flexibility. We designed a process that will pretreat the seawater to remove waste, colloids, bacteria, and membrane-damaging scalants. Once this seawater is treated is pumped to heavy pressures through a reverse osmosis system. After doing process optimization under various system parameter changes, we found a continuous, single stage, five element reverse osmosis process to be best. It produces a near 30% recovery. We reduce energy costs by including a pressure exchanger between our concentrate outlet and adjusted feed into the reverse osmosis system.

After performing mass and energy balances, we were able to size equipment and find our permeate composition. It turns out our process will meet L.A. standards for drinking water quality. However, post treatment will be done by the local LBWP to make the water taste better and be non-caustic to pipes. This will reduce “in-house” costs. Many environmental considerations were made, to meet our environmental standards. We meet EPA salinity requirements for our brine waste by diluting it. We also don’t use a lot of energy and thus reduce our fossil fuel consumption and CO2 generation. To be safe, we work well under max pressure limits for our equipment. We make sure we are ethical and meet all standards that our local county requires of us.

After an economic evaluation at a conservative well-water price of $700 per acre-foot, we realize that our revenue is way too low to cover our costs. Capital costs are especially large with regards to our large water tanks. We find that we reach a 10 year NPV breakpoint of $2104 per acre-foot after performing a sensitivity analysis. The Carlsbad Plant in San Diego worked out a deal for $2257 per acre-foot.

As we have met all design constraints for this project, we hope to convince policymakers to adopt our project at least $2104 per acre-foot water. If this deal is reached, we recommend that this project and design be funded. There is future work to be done to improve costs such as performing pump optimization. However, this work is minimal and won’t harm our bottom line if we reach our proposed price points.

__TOC__

=Introduction=

Due to a combination of various factors--including rising human population, climate change, increased energy consumption, and land erosion--the availability of fresh water as a resource for human consumption has been dwindling. This has had a negative impact on the world, but this impact has been especially felt in areas with natural shortages in fresh water. Desert regions, such as Saudi Arabia, have resorted to desalinating seawater at the coast and transporting up to their cities. This encompasses their entire freshwater system as very few natural water sources remain. In the United States there is a region which has shown increasing signs of following in the footsteps of Saudi Arabia and becoming fully reliant on desalination: Los Angeles, California.

=Location and Market Analysis=

In this project, Los Angeles was chosen as the location our plant, because the area has been subject to a severe drought for the last four years, as shown in Figure 1. The rapid decrease in the available water has caused permanent damages to the ground, decreased the amount of water in the reservoir from 45% to 25%, and reduced the amount of hydroelectric power produced in the state. Such deficiency in usable water has increased the water prices in Los Angeles by 6-7% annually for the last decade and is expected to rise by 3.4% annually for typical customers for the next five years. For customers who use more than 20000 gallons a year, the bill is expected to increase by 34% by 2021. This price increase is occurring because of both drought and maintenance. <ref>Pacific Institute: The California Drought. 2014.</ref>

[[File:Figure_1.jpg|right|200px|frame|Figure 1. California Drought Index.]]
As of March 2015 the Metropolitan Water District of Southern California was willing to pay as much as $700 per acre-foot of water to farmers in the Sacramento Valley, so we assume this will be the target price of the water produced by the this desalination plant.<ref>Hessel, Phil. “Dry Southern California Offers Northern Farmers Top Dollar for Water”. NBC News. March 18, 2015. </ref>

Currently, there is a plant in Carlsbad which supplies drinking water to the San Diego area. The plant can produce up to 50 million gallons per day and is the nation's largest water desalination plant. The water from this plant costs $2257 per acre-foot for the first 48000 acre-feet and about $2000 for any additional acre-foot. The county signed a contract with this plant to purchase the first 48000 acre-feet of water for the next 30 years.<ref>Carlsbad Desalination Plant. “Project Overview”. January 2016.</ref> Also, there is another desalination plant in Tampa that produces about 25 million gallons of water per day, and over 2000 other public and industrial plants with production capacity of over 300000 gallons per day. When conducting the economic analysis our plant, such numbers presented by other plants must be considered, so our project can yield positive profit.<ref>Leven, Rachel. “U.S. Desalination Industry Grows Since 2000; See as Essential to Meeting Supply Needs”. Bloomberg BNA. August 21, 2013.</ref>

The profitability of the project is naturally heavily dependent on the price of water in California. If the drought ends, water supply will increase and the water price will drop significantly, but if the drought continues, the price of water in California could rise even more rapidly and allow desalination plants to be more profitable.<ref>Southern California Public Radio. “Where is California water use decreasing?”. December 2015.</ref> It may be difficult to make this plant profitable without a guaranteed purchase or other subsidy from the local governments because the presence of pre-existing plants maintains the water price relatively low. However, external sources of profit such as guaranteed purchase or subsidy are expected because the ramifications of not having desalination plant or other sources of water may be significantly more expensive than the subsidizing a desalination plant assuming the drought continues.<ref>Weiser, Matt. “Could desalination solve California’s water problem?”. The Sacramento Bee. October 18, 2014.</ref>

=Design Basis and Technical Approach=

[[File:Figure_2.jpg|left|200px|frame|Figure 2. Vertical Well Feed]]
Considering the number of pre-existing desalination plant and the potential risk of the drought ending, our plant will be tentatively designed to produce 15 million gallons of drinkable water per day to homes and businesses in the Los Angeles county area. With a recovery rate of near 30%, and a requirement for diluting any brine, the input feed flowrate will be 70.5 million gallons per day. Following the similarly built Carlsbad plant, we expect the plant to last 30 years. The project building timeline is one year based on our relatively small plant size. We will run for 350 days a year while shutting down for cleaning and maintenance once every 3 months (based on DOW standards).<ref name = "Bates">Bates, Wayne. “Cleaning your RO”. Hydranautics. </ref> The plant will be built 500 meters from the coastline on Long Beach. A borehole will be drilled 10 meters into the ground. At this depth there should be seawater aquifers below the coast that regularly recharges itself with seawater from the ocean. A vertical well containing intake holes will be put into the borehole. The intake holes will be equipped with large size filters that can easily be replaced. At the surface, the well will be connected to a large-scale pump in a well pump house. From there, the seawater will be horizontally pumped into a seawater tank in the desalination plant. This will be considered the start of our process. The process ends after pretreatment, desalination, post treatment, and brine treatment are completed. Appendix 1 shows more details on the Design Basis.

The feed composition is assumed to be the standard seawater “average” based on the DOW Chemical ROSA 9.0 Technical Manual in Table 1. However, in reality the seawater composition at the plant site will have to be chemically tested. The final product will be based on L.A. standards for drinking water quality (Table 1). The specifications in the table are upper limits that we cannot exceed. We will aim to reach the salt level standards but any remineralization and pH treatment will be exported and paid for due to the added complexity and cost in making such a system in-house.

Due to the acidity of the desalinated water coming out of the reverse osmosis (RO) system (pH ~ 6), the pumps and pipes carrying the desalinated water for post treatment will likely corrode over time. We have decided to place the desalination plant by the Long Beach Water Reclamation Plant (about 50 meters away) to prevent pipe corrosion.

[[File:Table_1.jpg|center|200px|frame|Table 1. Concentration of the standard seawater, the output from the RO system, and LA standard.]]

=Desalination Overview and Alternatives=

We first began our research into the current landscape of the desalination industry, particularly of plants that were built in the United States. We discovered there were many processes but the main ones included: 
* Membrane Separation (Forward Osmosis and Reverse Osmosis)
* Thermal Separation (Multi Stage Flash Distillation)
* Phase Change Separation (Freeze Separation)
Each process was analyzed for their pros and cons. For our project design, we decided to go with reverse osmosis as our desalination process. This is because it is the most commonly used process in the industry, and for good reason. Unlike thermal separation, which requires huge amounts of energy to supply heat, reverse osmosis is very energy efficient. This reduces energy costs that would be spread across the plant’s lifetime.<ref name = "Republic">Republic of Mauritius NY. "Technical Aspects of Desalination Plant". Sustainable Sanitation and Water Management. </ref>

Forward osmosis is a newer technology being developed that also uses membranes to separate water and salt and may be more energy efficient overall than reverse osmosis. This is because forward osmosis doesn’t require a pump to pressurize seawater through the membranes, but rather leaves a higher salt concentrated solution (AKA draw solution) on the other side of the membrane for the seawater to be drawn into. Then a second step is required to regenerate the concentrated draw solution and produce purified water.<ref>Mallinson, Alissa. "Study Shows Forward Osmosis Desalination Not Energy Efficient." MIT News [Boston]. July 23, 2014.</ref> We decided to move away from forward osmosis because it is a process that is still being refined and developed for desalination and lacks the data, programs, and technologies that reverse osmosis users benefit from. Additionally, a recent study from MIT has claimed that due to the inherent nature of forward osmosis requiring that seawater be drawn to a higher concentrated salt solution before the water and salt is separated, it cannot possibly be more energy efficient than reverse osmosis.<ref>O'Hern, Sean C., Karnik, Rohit. Nanofiltration across Defect-Sealed Nanoporous Monolayer Graphene. American Chemical Society 15.5 (2014); 3254-260. </ref>

Other processes such as freeze separation were too restrictive in their feed and process demands and required a lot of capital. Ultimately, reverse osmosis is simply the best process that exists for desalination and has been ranked the best when performance is evaluated in terms of technical, environmental, and economic aspects. It is a process that is flexible in water quantity, operation start-up and shut-off, and site location.

Nevertheless there are disadvantages to reverse osmosis that must be considered. The feed seawater must not be too salty and too low quality. The pressure to pump the seawater is enormous. The construction time is long for large plants. The capital costs are also high when compared with other processes such as thermal separation. But all of these disadvantages are ultimately outweighed by the advantages reverse osmosis provides.<ref name = "Republic"/>

=Process Overview and Alternatives=

See Appendix 2 (Figure 1) for the Process Flow Diagram (PFD). The process itself is detailed below. The stream flow rates are in Appendix 2 (Table 1). We assumed 10 meters of pipe for every stream except the inlet and outlet streams of the entire PFD.

==Pretreatment==

After the 10 meter vertical well pumps seawater from the Long Beach coast, the feed is pumped 500 meters to a seawater tank that makes it easy to shut-down the system and reduce the large volume of flow. The level control keeps the seawater from overfilling the tank. From there, part of the seawater goes to brine treatment while the rest go to pretreatment. The purpose of pretreatment in the desalination process is to prevent equipment damage by removing particles and microorganisms from the water. A variety of pollutants -- including calcium carbonate, sulfates, silica, clay, bacteria, and other colloids and microorganisms -- can disrupt the reverse osmosis process. Media filtration uses fine grains such as sand or anthracite to filter the particles. We used a large filter to catch larger particle material (e.g. sand). However, simple filtration is not enough to prevent colloid fouling. An ultrafilter (i.e. microfiltration) is used to remove virtually all suspended matter, effectively preventing colloid fouling. The costs associated with ultrafiltration are not prohibitive in terms of the project’s budget.<ref name = "Desalination">Desalination Pre-treatment. Lenntech.</ref>

Reverse osmosis systems can also be disrupted by the growth of microbiological films of bacteria on the film and pipe walls. They are typically treated as colloids during pretreatment, and most organisms are removed from the water by ultrafiltration, but this does not remove the smallest of organic particles. DBNPA is the most effective biocide, but these aren’t often used and have to be removed from the water. Chlorination is an effective way to remove bacteria, but it can cause degradation of the membrane, and must be removed before the reverse osmosis occurs. In our process we will use UV irradiation to remove any organic material that passes through the microfilters.<ref name = "Desalination"/>

Scaling can reduce the flux across the membrane and can build up in pipe corners, causing backup. It is caused by inorganic salts -- specifically sulfates and calcium and magnesium carbonate -- present in the water. Scaling can be prevented and/or controlled with acids, antiscalants, or ion exchange softening. Antiscalants are good for removing sulfates, but not much else. Acidification will only be effective at removing CaCO3 from water but, the decreased pH may deteriorate the membrane and pipes, so it may require future research.<ref name = "Desalination"/> Ion exchange is very expensive and requires more capital to undergo, but is overall the more effective option.

The ion exchanger is a pretreatment system which acts as a method to fight membrane scaling in reverse osmosis. The ion exchanger works by using sodium carboxylate as a resin which will react with Chlorine and Magnesium ions, removing them from the feed. The sodium is substituted out and enters the feed stream, increasing the overall sodium level and TDS. The lack of magnesium and chlorine ions means that there is nothing to react with the carbonates to form scaling salts on the membranes.

==Desalination (Reverse Osmosis)==

Two pumps (P-105 & P-106) take the adjusted feed seawater up to 58 bar. However, before this pumping occurs, about 70% of the original stream gets separated and heads into a PX Pressure Exchanger. This separated stream later joins together with the pumped stream to enter the reverse osmosis system at the required pressure of 58 bar. The streams can combine together as they have equal pressure. It is the concentrate stream from the output of the reverse osmosis system that transfers its energy to the adjusted feed in the PX Pressure Exchanger.

[[File:Figure_3.jpg|left|300px|frame|Figure 3. Inside look of pressure exchanger]]
The pressure of the concentrate stream is about 57 bar and it is fed into an exchanger which works by using vessel that resembles a small pipe within a rotating wheel. The low pressure (LP) feed enters the vessel at 1 bar (see Figure 3). The feed turns and is pushed out into an outlet pipe by the high pressure (HP) concentrate stream at ~57 bar. The feed is now a HP stream via this exchanged energy. The now LP concentrate stream turns and is pushed out into an outlet pipe by the LP feed entering the vessel.

98% of the energy within the concentrate stream can be recovered via the PX Pressure Exchanger.<ref>Desalination Product and Applications Catalog 2013. “Advance Technologies for Seawater Reverse Osmosis and Brackish Water Applications”. Energy Recovery, Inc. 2012-2014.</ref >Thus, the now HP feed stream is not quite 58 bar and will need to be pumped by a small booster pump (P-107).

When designing a reverse osmosis system, we need to make decisions on what specific kind of reverse osmosis system will be implemented. When making the decision, we investigated the advantages of different types of reverse osmosis among the following options: batch vs. continuous process, number of stages, and number of elements.<ref>FILMTEC Reverse Osmosis Membranes. “Water & Process Solutions Technical Manual.” Dow Chemical Company.</ref> Our decision was made based on process optimization which is described later in this report. The end result of that optimization was to go with a single stage, continuous process with five elements in series.

=Post Treatment=

The permeate is sent to post treatment after the salt content is analyzed and controlled if too high or too low due to any unexpected errors. As stated above, the post treatment process will be done 50 meters from our plant by the Long Beach Water Reclamation Plant. The reason we aren’t performing post treatment “in-house” is due to the increased cost it would take for us to build and run the equipment for post treatment. We have the option to use the government’s water treatment facilities that already exist and our 15 million gallons per day output is a relatively small addition to the large Long Beach Water Reclamation Plant. We will assume any costs to post treatment will be pushed to the consumer in the price of purchasing the drinking water from the Long Beach Water Reclamation Plant.

Post treatment at any facility works to bring water to “taste standards”. The water that is the permeate of our process is already meeting L.A. drinking standards as shown above in Table 1 after our mass balance calculations around the reverse osmosis system (Appendix 3, Table 1). The post treatment increases the pH from 6.3 to 7.3 with the addition of base and then adds minerals like fluoride, calcium, and magnesium back in to improve taste.

==Brine Treatment and Byproduct Waste==

A natural consequence of a reverse osmosis procedure is that a lot of salt will be leftover from the membrane separation in a very concentrated form. This is called “brine” and will have to be taken care of. According to EPA requirements, the salt concentration of water entering the environment (via land or sea) cannot be higher than 40,000 ppm.8 The salt concentration of our brine is much higher (around 70,000 ppm). Thus, the brine must be removed or treated. We decided to treat the brine by diluting it as removing brine can be difficult and may require lots of land (e.g. evaporation ponds).

After the concentrate exchanges energy in the energy exchanger, it becomes a low pressure stream. It is pumped into a dilution tank and mixes with 18.2 million gallons per day of seawater from the inlet to become diluted brine which is pumped 600 meters out to sea.

The waste byproducts are the solid buildup on the filters and membranes. The filters and membranes also have a set lifetime and thus will also need to be removed. The waste is all non-hazardous and thus can be put in the municipal dump.

=Process Details and Equipment Sizing=

==Filters and UV Treatment [Pretreatment]==
The cost of our ultrafiltration system is estimated using data from the University of Tennessee at Knoxville.<ref>Qiang He and R. Bruce Robinson Membrane Filtration Systems. “Membrane Module Costs.” University of Tennessee at Knoxville. December 2014.</ref> It estimated the cost of the cost of the ultrafiltration $0.057 per 1000 gallons and the construction and capital costs are $0.39 per gallon per day. The cost of UV treatment is based off of units available from buyultraviolet.com for units that can treat about 450 gallons per minute.<ref>Atlantic Ultraviolet Corporation. “Megatron Water Disinfection System.”</ref> Due to the water having to be in relatively small pipes for treatment we do not expect the costs to scale very much, so we believe these units to be an accurate cost estimate.

==Ion Exchanger [Pretreatment] ==

The ion exchanger was sized using data from DOW. Our process will use 60 vessels with ion exchange beds of volume 5.73 m3. With a flow rate of 30 bed volumes/hour (10,312.2 m3/hr) the system of beds has a capacity of 125% of the feed. This allows for periodic regeneration of the resin, a weak acid cation resin (DOW MAC-3 Resin). This weak cation resin will remove cationic calcium and magnesium from the water to prevent scaling. The mass flow was calculated via a mass balance on the ion exchanger (Appendix 3).

==Reverse Osmosis System [Desalination]==

Each of the 2607 Reverse Osmosis SW30HRLE-440i [Seawater 30 High Rejection Low Energy 440 ft2 interlocking cap membrane] units have the same dimensions. A, the net length, is 1016 mm. B, the total length including the endcap, is 1029 mm. C, the diameter of the filter element, is 201mm. D, the diameter of the permeate pathway, is 29 mm. For a single stage, we have five of the elements in series so the overall length of a single stage is expected to be roughly 5145 mm. The membrane material is polyamide thin film composite. It will last three years.
Operating Conditions [ROSA 9.0]: 
* Temperature: 30oC 
* Pressure: 58 bar 
* Flow Rate: 52.3 Million gal per day 
* Recovery: 28.7%

[[File:Figure_4.png|center|400px|frame|Figure 4. Diagram of Dow FILMTEC SW30HRLE-440i]]
Cleaning Schedule: 
* Conservative Estimate: Once every three months (three days available for cleaning)<ref name = "Bates"/> 
* DOW FILMTEC Cleaning Conditions 
** The normalized permeate flow drops 10%
* The normalized salt passage increases 5 - 10%

==Seawater Holding Tank and Dilution Tank [Brine Treatment]==

The initial seawater holding tank works as a buffer to prevent surges and help perform shutdowns quickly and easily. We expect to run continuously for 350 days a year with the rest of the available time spread out and used for shutting down the plant for cleaning membranes, filters, and the fixing of unexpected errors. There is a big tradeoff between having a holding tank with high retention time and having a holding tank with low cost. With that in mind, we erred on the side of lowering capital cost and went with a retention time of one hour. The volume could be found by multiplying the retention time with the feed flow rate (Stream 1). We assumed that a good height to width ratio for a large water tank would be 4:1.

A mass balance was done around the dilution tank to find what amount of seawater would be necessary to dilute the brine to 45,000 ppm salinity (Appendix 3; Table 3). With this the feed flow rate into the tank could be found. It was found that it takes about 20 minutes of retention time for good dilution of the brine in seawater.8 We again assumed that a good height to width ratio for a large water tank would be 4:1.

We assumed a corrosion allowance of 4 mm wall thickness. We calculated wall thickness using vessel pressure, hoop stress, longitudinal stress, and weld factor. Each tank is made with SS316 to protect against corrosion.<ref>Towler G, Sinnott R. Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013.</ref>

<center>
'''Table 2.''' Tank sizes.
[[File:Table_2.jpg|center|600px]]
</center>

=Process Optimization=
<center> '''Table 3.'''Optimization Criteria
[[File:Table_3.jpg|center|400px]]
</center>
In order to optimize the reverse osmosis system, numerous simulations were run. Three representative cases are reported as shown above (Table 3). The objective of the optimization was to minimize the required water price to break even in 10 years. In order to achieve this goal, the balance between the capital cost from the number of the filter elements and utility cost from energy consumption was weighed. Before the energy recovery system was implemented, the general trend was that if we have more filter elements in series, overall recovery increases but recovery rate per filter decreases, and we need more filter elements. While increasing the elements in series increase the required number of filters, and thereby increases the capital cost, we require less energy to produce the same amount of water. After the energy recovery system is implemented, the correlation between the number of elements in series and the required energy becomes more complicated because high pressure concentrate is used to pressurize the feed. However, the trade-off between the utility cost and capital cost still exists. Of the three cases shown above, single stage, five elements RO system turned out to be the most optimal solution because single stage, single element system required higher utility cost and two stage system slightly did not meet the LA standard. Detailed information which was used for the optimization is available in Appendix 3 and in Appendix 6.

=Ethical, Environmental, and Societal Considerations=

There are a few key considerations we have followed to meet ethical, environmental, and societal guidelines and requirements: 
* Meet L.A. Drinking Water Standards
** Correct ion concentration
* Moved near water treatment plant
** Pipes will not corrode from low pH
* Working below pressure limits
* Water will be locally served and meet taste standards
** Via Post Treatment mineralization
* Follows EPA Environmental Salinity Requirements
* Reduced Energy Requirements
** PX Pressure Exchanger
** Optimized RO System
** Reverse Osmosis is itself Energy Saving

For further environmental considerations, see effluent summary in Appendix 4.

=Economic Analysis=

The pricing of the majority of the process was completed using the Aspen Economic Analyzer, however we had determine costs for the ultrafiltration, reverse osmosis membranes, and the ultraviolet treatment was completed separately. This led to a total ISBL of $56 MM, OSBL of $22.4 MM, Engineering Costs of $22.4 MM and a Contingency of $7.8 MM. We also chose a working Capital of $10 MM. This working capital is sufficient because we have no raw materials costs. We also had FCOP of $5.2 MM and VCOP of $3.5 MM.

We originally were planning on pricing our water at $700 per acre-foot, but after doing economic analysis we had a 10 year NPV of $-76.2 MM. After determining did a sensitivity analysis and determined that if our water was priced at $2104 per acre-foot we could have an NPV of $0 after 10 years and if we were able to get the same price of $2257 per acre-foot as the plant in Carlsbad we would have a 10 year NPV of $8.3 MM which increases to $41.1 MM at 20 years. For more detailed data about the economic analysis see Appendix 6.

=Conclusion and Recommendations=

Our mission was to create an economically viable process that will desalinate water for human consumption in the L.A. County. We have found after research and optimization that the best process to undergo desalination is a continuous single stage, five elements reverse osmosis system. Our process can meet the required L.A. drinking water standard and we have set up plans for post treatment at a local water reclamation site. While meeting the drinking water standards, we also met our ethical, environmental, and social standards.

Nevertheless, at a sale price of $700 per acre-foot, our proposed desalination plant in Los Angeles, California looks to be a very expensive project to undergo. Without any sort of government assistance to carry the heavy capital costs of such a project, we would suggest the plant not be built. However, after performing a sensitivity analysis at higher (and less conservative) price points we reached a breakeven point in 10 years with a sale price of $2104 per acre-foot of drinking water. Research has shown that the droughts striking California look only to get worse in the near future. With limited freshwater resources, there is no option other than to consider desalination. Therefore it is reasonable to expect we could get near the same price deal as local desalination plants like the Carlsbad Plant ($2257 per acre-foot) As we have met all design constraints for this project, we hope to convince policymakers to adopt our project at least $2104 per acre-foot water. If this deal is reached, we recommend that this project and design be funded.

However, there are ways to improve our design with future work. Right now all the pipes diameters are based on having a flowrate of 1.5 m/s. But in reality, we’d need to use standard piping sizes to get more accurate pricing and energy costs. Also, in reality we would be purchasing some of these vessels in bulk quantities (i.e. bulk filters) so the price could be lowered with negotiation and economies of scale. We also have pumps that are providing very little pressure for their high cost. We can remove these pumps and simply run at higher pressures earlier. We can potentially reduce early energy costs by building a “floating” desalination plant in the sea to reduce pumping distance from the well.

=Appendices=

==Appendix 1: Design Basis==
[[File:Appendix_1.png|thumb|border|center|800px|middle|]]

==Appendix 2: Process Flow Diagram==
[[File:Appendix_2.png|center]]
==Appendix 3: Mass and Energy Balance Tables==
'''Table 1'''. Mass balance on 1 stage 5 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T1.png|center|600px|]]

'''Table 2'''. Mass balance on 1 stage 1 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T2.png|center|600px|]]

'''Table 3'''. Mass balance on 2 stage 5,5 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T3.png|center|600px|]]

'''Table 4'''. 1 stage 5 elements CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T4.png|center|600px|]]

'''Table 5'''. 1 stage 1 element CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T5.png|center|600px|]]

'''Table 6'''. 2 stages 5,5 elements CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T6.png|center|600px|]]

'''Table 7'''. Energy balance on 1 stage 5 elements RO system
[[File:Appendix_3_T7.png|center|600px|]]

==Appendix 4: Effluent Summary==
[[File:Appendix_4.png|center]]

==Appendix 5: Cost of Production==
[[File:Appendix_5.png|center]]

==Appendix 6: Economic Anlaysis==
===Case I ===
Water is $700; RO with 1 stage and 5 elements
[[File:Appendix_6_C1.png|center]]

===Case II===
Water is $1400; RO with 1 stage and 5 elements
[[File:Appendix_6_C2.png|center]]

===Case III: Carlsbad ===
Water is $2257; RO with 1 stage and 5 elements
[[File:Appendix_6_C3.png|center]]

===Case IV: Breakeven (NPV = 0 at 10 years)===
Water is $2104; RO with 1 stage and 5 elements
[[File:Appendix_6_C4.png|center]]

===Case V: Breakeven (NPV = 0 at 10 years)===
Water is $2112; RO with 2 stages and 5,5 elements
[[File:Appendix_6_C5.png|center]]

===Case VI: Breakeven (NPV = 0 at 10 years)===
Water is $2138 with 1 stage and 1 element
[[File:Appendix_6_C6.png|center]]

==Appendix 7: ROSA Simulation Outputs==

'''Table 1'''. System detail of 1 stage 5 elements RO system
[[File:Appendix_7_T1.png|center]]

'''Table 2'''. System detail of 1 stage 1 elements RO system
[[File:Appendix_7_T2.png|center]]

'''Table 3'''. System detail of 2 stage 5,5 elements system
[[File:Appendix_7_T3.png|center]]

'''Table 4'''. 1 stage 5 elements RO system concentrations and pH in each stream
[[File:Appendix_7_T4.png|center]]

'''Table 5'''. 1 stage 1 element RO system concentration and pH in each stream
[[File:Appendix_7_T5.png|center]]

'''Table 6'''. 2 stage 5,5 elements RO system concentrations and pH in each stream
[[File:Appendix_7_T6.png|center]]

=References=
<references/>

Desalination - Team E

2016-03-12T03:30:07Z

Bts511: /* Appendix 1: Design Basis */

Group E Corporation

Authors: Hassan Ali, Woo Soo Choe, Brett Sleyster, Jake Stolley

Instructors: Fengqi You, David Wegerer

March 11, 2016

=Executive Summary=

Outlined in this report is the project model report for the proposed desalination plant in Los Angeles, California which will provide drinking water to local citizens using reverse osmosis process. This location was chosen due to the region’s propensity for droughts in recent years. The fresh water resources in California have plummeted as water prices have risen 3-4% each year. This is only expected to continue. Other companies have worked deals with the Californian government to create desalination plants that provide a new fresh water drinking source. This suggests that there is a strong market here with potential growth.

The plant will have an output of 15 million gallons per day, which is about medium size relative to other plants. Thus the build time will be about one year. The plant will run for 30 years, with 350 days per year of use. Shutdown time will be used for cleaning and maintenance of equipment. The plant will be built on Long Beach, 500 meters from the coast nearby the Long Beach Water Reclamation Plant (LBWP). The input will withdraw 70.5 million gallons per day of seawater that will be used for reverse osmosis and dilution of brine.

Reverse osmosis was chosen as the desalination method for its low energy cost and flexibility. We designed a process that will pretreat the seawater to remove waste, colloids, bacteria, and membrane-damaging scalants. Once this seawater is treated is pumped to heavy pressures through a reverse osmosis system. After doing process optimization under various system parameter changes, we found a continuous, single stage, five element reverse osmosis process to be best. It produces a near 30% recovery. We reduce energy costs by including a pressure exchanger between our concentrate outlet and adjusted feed into the reverse osmosis system.

After performing mass and energy balances, we were able to size equipment and find our permeate composition. It turns out our process will meet L.A. standards for drinking water quality. However, post treatment will be done by the local LBWP to make the water taste better and be non-caustic to pipes. This will reduce “in-house” costs. Many environmental considerations were made, to meet our environmental standards. We meet EPA salinity requirements for our brine waste by diluting it. We also don’t use a lot of energy and thus reduce our fossil fuel consumption and CO2 generation. To be safe, we work well under max pressure limits for our equipment. We make sure we are ethical and meet all standards that our local county requires of us.

After an economic evaluation at a conservative well-water price of $700 per acre-foot, we realize that our revenue is way too low to cover our costs. Capital costs are especially large with regards to our large water tanks. We find that we reach a 10 year NPV breakpoint of $2104 per acre-foot after performing a sensitivity analysis. The Carlsbad Plant in San Diego worked out a deal for $2257 per acre-foot.

As we have met all design constraints for this project, we hope to convince policymakers to adopt our project at least $2104 per acre-foot water. If this deal is reached, we recommend that this project and design be funded. There is future work to be done to improve costs such as performing pump optimization. However, this work is minimal and won’t harm our bottom line if we reach our proposed price points.

__TOC__

=Introduction=

Due to a combination of various factors--including rising human population, climate change, increased energy consumption, and land erosion--the availability of fresh water as a resource for human consumption has been dwindling. This has had a negative impact on the world, but this impact has been especially felt in areas with natural shortages in fresh water. Desert regions, such as Saudi Arabia, have resorted to desalinating seawater at the coast and transporting up to their cities. This encompasses their entire freshwater system as very few natural water sources remain. In the United States there is a region which has shown increasing signs of following in the footsteps of Saudi Arabia and becoming fully reliant on desalination: Los Angeles, California.

=Location and Market Analysis=

In this project, Los Angeles was chosen as the location our plant, because the area has been subject to a severe drought for the last four years, as shown in Figure 1. The rapid decrease in the available water has caused permanent damages to the ground, decreased the amount of water in the reservoir from 45% to 25%, and reduced the amount of hydroelectric power produced in the state. Such deficiency in usable water has increased the water prices in Los Angeles by 6-7% annually for the last decade and is expected to rise by 3.4% annually for typical customers for the next five years. For customers who use more than 20000 gallons a year, the bill is expected to increase by 34% by 2021. This price increase is occurring because of both drought and maintenance. <ref>Pacific Institute: The California Drought. 2014.</ref>

[[File:Figure_1.jpg|right|200px|frame|Figure 1. California Drought Index.]]
As of March 2015 the Metropolitan Water District of Southern California was willing to pay as much as $700 per acre-foot of water to farmers in the Sacramento Valley, so we assume this will be the target price of the water produced by the this desalination plant.<ref>Hessel, Phil. “Dry Southern California Offers Northern Farmers Top Dollar for Water”. NBC News. March 18, 2015. </ref>

Currently, there is a plant in Carlsbad which supplies drinking water to the San Diego area. The plant can produce up to 50 million gallons per day and is the nation's largest water desalination plant. The water from this plant costs $2257 per acre-foot for the first 48000 acre-feet and about $2000 for any additional acre-foot. The county signed a contract with this plant to purchase the first 48000 acre-feet of water for the next 30 years.<ref>Carlsbad Desalination Plant. “Project Overview”. January 2016.</ref> Also, there is another desalination plant in Tampa that produces about 25 million gallons of water per day, and over 2000 other public and industrial plants with production capacity of over 300000 gallons per day. When conducting the economic analysis our plant, such numbers presented by other plants must be considered, so our project can yield positive profit.<ref>Leven, Rachel. “U.S. Desalination Industry Grows Since 2000; See as Essential to Meeting Supply Needs”. Bloomberg BNA. August 21, 2013.</ref>

The profitability of the project is naturally heavily dependent on the price of water in California. If the drought ends, water supply will increase and the water price will drop significantly, but if the drought continues, the price of water in California could rise even more rapidly and allow desalination plants to be more profitable.<ref>Southern California Public Radio. “Where is California water use decreasing?”. December 2015.</ref> It may be difficult to make this plant profitable without a guaranteed purchase or other subsidy from the local governments because the presence of pre-existing plants maintains the water price relatively low. However, external sources of profit such as guaranteed purchase or subsidy are expected because the ramifications of not having desalination plant or other sources of water may be significantly more expensive than the subsidizing a desalination plant assuming the drought continues.<ref>Weiser, Matt. “Could desalination solve California’s water problem?”. The Sacramento Bee. October 18, 2014.</ref>

=Design Basis and Technical Approach=

[[File:Figure_2.jpg|left|200px|frame|Figure 2. Vertical Well Feed]]
Considering the number of pre-existing desalination plant and the potential risk of the drought ending, our plant will be tentatively designed to produce 15 million gallons of drinkable water per day to homes and businesses in the Los Angeles county area. With a recovery rate of near 30%, and a requirement for diluting any brine, the input feed flowrate will be 70.5 million gallons per day. Following the similarly built Carlsbad plant, we expect the plant to last 30 years. The project building timeline is one year based on our relatively small plant size. We will run for 350 days a year while shutting down for cleaning and maintenance once every 3 months (based on DOW standards).<ref name = "Bates">Bates, Wayne. “Cleaning your RO”. Hydranautics. </ref> The plant will be built 500 meters from the coastline on Long Beach. A borehole will be drilled 10 meters into the ground. At this depth there should be seawater aquifers below the coast that regularly recharges itself with seawater from the ocean. A vertical well containing intake holes will be put into the borehole. The intake holes will be equipped with large size filters that can easily be replaced. At the surface, the well will be connected to a large-scale pump in a well pump house. From there, the seawater will be horizontally pumped into a seawater tank in the desalination plant. This will be considered the start of our process. The process ends after pretreatment, desalination, post treatment, and brine treatment are completed. Appendix 1 shows more details on the Design Basis.

The feed composition is assumed to be the standard seawater “average” based on the DOW Chemical ROSA 9.0 Technical Manual in Table 1. However, in reality the seawater composition at the plant site will have to be chemically tested. The final product will be based on L.A. standards for drinking water quality (Table 1). The specifications in the table are upper limits that we cannot exceed. We will aim to reach the salt level standards but any remineralization and pH treatment will be exported and paid for due to the added complexity and cost in making such a system in-house.

Due to the acidity of the desalinated water coming out of the reverse osmosis (RO) system (pH ~ 6), the pumps and pipes carrying the desalinated water for post treatment will likely corrode over time. We have decided to place the desalination plant by the Long Beach Water Reclamation Plant (about 50 meters away) to prevent pipe corrosion.

[[File:Table_1.jpg|center|200px|frame|Table 1. Concentration of the standard seawater, the output from the RO system, and LA standard.]]

=Desalination Overview and Alternatives=

We first began our research into the current landscape of the desalination industry, particularly of plants that were built in the United States. We discovered there were many processes but the main ones included: 
* Membrane Separation (Forward Osmosis and Reverse Osmosis)
* Thermal Separation (Multi Stage Flash Distillation)
* Phase Change Separation (Freeze Separation)
Each process was analyzed for their pros and cons. For our project design, we decided to go with reverse osmosis as our desalination process. This is because it is the most commonly used process in the industry, and for good reason. Unlike thermal separation, which requires huge amounts of energy to supply heat, reverse osmosis is very energy efficient. This reduces energy costs that would be spread across the plant’s lifetime.<ref name = "Republic">Republic of Mauritius NY. "Technical Aspects of Desalination Plant". Sustainable Sanitation and Water Management. </ref>

Forward osmosis is a newer technology being developed that also uses membranes to separate water and salt and may be more energy efficient overall than reverse osmosis. This is because forward osmosis doesn’t require a pump to pressurize seawater through the membranes, but rather leaves a higher salt concentrated solution (AKA draw solution) on the other side of the membrane for the seawater to be drawn into. Then a second step is required to regenerate the concentrated draw solution and produce purified water.<ref>Mallinson, Alissa. "Study Shows Forward Osmosis Desalination Not Energy Efficient." MIT News [Boston]. July 23, 2014.</ref> We decided to move away from forward osmosis because it is a process that is still being refined and developed for desalination and lacks the data, programs, and technologies that reverse osmosis users benefit from. Additionally, a recent study from MIT has claimed that due to the inherent nature of forward osmosis requiring that seawater be drawn to a higher concentrated salt solution before the water and salt is separated, it cannot possibly be more energy efficient than reverse osmosis.<ref>O'Hern, Sean C., Karnik, Rohit. Nanofiltration across Defect-Sealed Nanoporous Monolayer Graphene. American Chemical Society 15.5 (2014); 3254-260. </ref>

Other processes such as freeze separation were too restrictive in their feed and process demands and required a lot of capital. Ultimately, reverse osmosis is simply the best process that exists for desalination and has been ranked the best when performance is evaluated in terms of technical, environmental, and economic aspects. It is a process that is flexible in water quantity, operation start-up and shut-off, and site location.

Nevertheless there are disadvantages to reverse osmosis that must be considered. The feed seawater must not be too salty and too low quality. The pressure to pump the seawater is enormous. The construction time is long for large plants. The capital costs are also high when compared with other processes such as thermal separation. But all of these disadvantages are ultimately outweighed by the advantages reverse osmosis provides.<ref name = "Republic"/>

=Process Overview and Alternatives=

See Appendix 2 (Figure 1) for the Process Flow Diagram (PFD). The process itself is detailed below. The stream flow rates are in Appendix 2 (Table 1). We assumed 10 meters of pipe for every stream except the inlet and outlet streams of the entire PFD.

==Pretreatment==

After the 10 meter vertical well pumps seawater from the Long Beach coast, the feed is pumped 500 meters to a seawater tank that makes it easy to shut-down the system and reduce the large volume of flow. The level control keeps the seawater from overfilling the tank. From there, part of the seawater goes to brine treatment while the rest go to pretreatment. The purpose of pretreatment in the desalination process is to prevent equipment damage by removing particles and microorganisms from the water. A variety of pollutants -- including calcium carbonate, sulfates, silica, clay, bacteria, and other colloids and microorganisms -- can disrupt the reverse osmosis process. Media filtration uses fine grains such as sand or anthracite to filter the particles. We used a large filter to catch larger particle material (e.g. sand). However, simple filtration is not enough to prevent colloid fouling. An ultrafilter (i.e. microfiltration) is used to remove virtually all suspended matter, effectively preventing colloid fouling. The costs associated with ultrafiltration are not prohibitive in terms of the project’s budget.<ref name = "Desalination">Desalination Pre-treatment. Lenntech.</ref>

Reverse osmosis systems can also be disrupted by the growth of microbiological films of bacteria on the film and pipe walls. They are typically treated as colloids during pretreatment, and most organisms are removed from the water by ultrafiltration, but this does not remove the smallest of organic particles. DBNPA is the most effective biocide, but these aren’t often used and have to be removed from the water. Chlorination is an effective way to remove bacteria, but it can cause degradation of the membrane, and must be removed before the reverse osmosis occurs. In our process we will use UV irradiation to remove any organic material that passes through the microfilters.<ref name = "Desalination"/>

Scaling can reduce the flux across the membrane and can build up in pipe corners, causing backup. It is caused by inorganic salts -- specifically sulfates and calcium and magnesium carbonate -- present in the water. Scaling can be prevented and/or controlled with acids, antiscalants, or ion exchange softening. Antiscalants are good for removing sulfates, but not much else. Acidification will only be effective at removing CaCO3 from water but, the decreased pH may deteriorate the membrane and pipes, so it may require future research.<ref name = "Desalination"/> Ion exchange is very expensive and requires more capital to undergo, but is overall the more effective option.

The ion exchanger is a pretreatment system which acts as a method to fight membrane scaling in reverse osmosis. The ion exchanger works by using sodium carboxylate as a resin which will react with Chlorine and Magnesium ions, removing them from the feed. The sodium is substituted out and enters the feed stream, increasing the overall sodium level and TDS. The lack of magnesium and chlorine ions means that there is nothing to react with the carbonates to form scaling salts on the membranes.

==Desalination (Reverse Osmosis)==

Two pumps (P-105 & P-106) take the adjusted feed seawater up to 58 bar. However, before this pumping occurs, about 70% of the original stream gets separated and heads into a PX Pressure Exchanger. This separated stream later joins together with the pumped stream to enter the reverse osmosis system at the required pressure of 58 bar. The streams can combine together as they have equal pressure. It is the concentrate stream from the output of the reverse osmosis system that transfers its energy to the adjusted feed in the PX Pressure Exchanger.

[[File:Figure_3.jpg|left|300px|frame|Figure 3. Inside look of pressure exchanger]]
The pressure of the concentrate stream is about 57 bar and it is fed into an exchanger which works by using vessel that resembles a small pipe within a rotating wheel. The low pressure (LP) feed enters the vessel at 1 bar (see Figure 3). The feed turns and is pushed out into an outlet pipe by the high pressure (HP) concentrate stream at ~57 bar. The feed is now a HP stream via this exchanged energy. The now LP concentrate stream turns and is pushed out into an outlet pipe by the LP feed entering the vessel.

98% of the energy within the concentrate stream can be recovered via the PX Pressure Exchanger.<ref>Desalination Product and Applications Catalog 2013. “Advance Technologies for Seawater Reverse Osmosis and Brackish Water Applications”. Energy Recovery, Inc. 2012-2014.</ref >Thus, the now HP feed stream is not quite 58 bar and will need to be pumped by a small booster pump (P-107).

When designing a reverse osmosis system, we need to make decisions on what specific kind of reverse osmosis system will be implemented. When making the decision, we investigated the advantages of different types of reverse osmosis among the following options: batch vs. continuous process, number of stages, and number of elements.<ref>FILMTEC Reverse Osmosis Membranes. “Water & Process Solutions Technical Manual.” Dow Chemical Company.</ref> Our decision was made based on process optimization which is described later in this report. The end result of that optimization was to go with a single stage, continuous process with five elements in series.

=Post Treatment=

The permeate is sent to post treatment after the salt content is analyzed and controlled if too high or too low due to any unexpected errors. As stated above, the post treatment process will be done 50 meters from our plant by the Long Beach Water Reclamation Plant. The reason we aren’t performing post treatment “in-house” is due to the increased cost it would take for us to build and run the equipment for post treatment. We have the option to use the government’s water treatment facilities that already exist and our 15 million gallons per day output is a relatively small addition to the large Long Beach Water Reclamation Plant. We will assume any costs to post treatment will be pushed to the consumer in the price of purchasing the drinking water from the Long Beach Water Reclamation Plant.

Post treatment at any facility works to bring water to “taste standards”. The water that is the permeate of our process is already meeting L.A. drinking standards as shown above in Table 1 after our mass balance calculations around the reverse osmosis system (Appendix 3, Table 1). The post treatment increases the pH from 6.3 to 7.3 with the addition of base and then adds minerals like fluoride, calcium, and magnesium back in to improve taste.

==Brine Treatment and Byproduct Waste==

A natural consequence of a reverse osmosis procedure is that a lot of salt will be leftover from the membrane separation in a very concentrated form. This is called “brine” and will have to be taken care of. According to EPA requirements, the salt concentration of water entering the environment (via land or sea) cannot be higher than 40,000 ppm.8 The salt concentration of our brine is much higher (around 70,000 ppm). Thus, the brine must be removed or treated. We decided to treat the brine by diluting it as removing brine can be difficult and may require lots of land (e.g. evaporation ponds).

After the concentrate exchanges energy in the energy exchanger, it becomes a low pressure stream. It is pumped into a dilution tank and mixes with 18.2 million gallons per day of seawater from the inlet to become diluted brine which is pumped 600 meters out to sea.

The waste byproducts are the solid buildup on the filters and membranes. The filters and membranes also have a set lifetime and thus will also need to be removed. The waste is all non-hazardous and thus can be put in the municipal dump.

=Process Details and Equipment Sizing=

==Filters and UV Treatment [Pretreatment]==
The cost of our ultrafiltration system is estimated using data from the University of Tennessee at Knoxville.<ref>Qiang He and R. Bruce Robinson Membrane Filtration Systems. “Membrane Module Costs.” University of Tennessee at Knoxville. December 2014.</ref> It estimated the cost of the cost of the ultrafiltration $0.057 per 1000 gallons and the construction and capital costs are $0.39 per gallon per day. The cost of UV treatment is based off of units available from buyultraviolet.com for units that can treat about 450 gallons per minute.<ref>Atlantic Ultraviolet Corporation. “Megatron Water Disinfection System.”</ref> Due to the water having to be in relatively small pipes for treatment we do not expect the costs to scale very much, so we believe these units to be an accurate cost estimate.

==Ion Exchanger [Pretreatment] ==

The ion exchanger was sized using data from DOW. Our process will use 60 vessels with ion exchange beds of volume 5.73 m3. With a flow rate of 30 bed volumes/hour (10,312.2 m3/hr) the system of beds has a capacity of 125% of the feed. This allows for periodic regeneration of the resin, a weak acid cation resin (DOW MAC-3 Resin). This weak cation resin will remove cationic calcium and magnesium from the water to prevent scaling. The mass flow was calculated via a mass balance on the ion exchanger (Appendix 3).

==Reverse Osmosis System [Desalination]==

Each of the 2607 Reverse Osmosis SW30HRLE-440i [Seawater 30 High Rejection Low Energy 440 ft2 interlocking cap membrane] units have the same dimensions. A, the net length, is 1016 mm. B, the total length including the endcap, is 1029 mm. C, the diameter of the filter element, is 201mm. D, the diameter of the permeate pathway, is 29 mm. For a single stage, we have five of the elements in series so the overall length of a single stage is expected to be roughly 5145 mm. The membrane material is polyamide thin film composite. It will last three years.
Operating Conditions [ROSA 9.0]: 
* Temperature: 30oC 
* Pressure: 58 bar 
* Flow Rate: 52.3 Million gal per day 
* Recovery: 28.7%

[[File:Figure_4.png|center|400px|frame|Figure 4. Diagram of Dow FILMTEC SW30HRLE-440i]]
Cleaning Schedule: 
* Conservative Estimate: Once every three months (three days available for cleaning)<ref name = "Bates"/> 
* DOW FILMTEC Cleaning Conditions 
** The normalized permeate flow drops 10%
* The normalized salt passage increases 5 - 10%

==Seawater Holding Tank and Dilution Tank [Brine Treatment]==

The initial seawater holding tank works as a buffer to prevent surges and help perform shutdowns quickly and easily. We expect to run continuously for 350 days a year with the rest of the available time spread out and used for shutting down the plant for cleaning membranes, filters, and the fixing of unexpected errors. There is a big tradeoff between having a holding tank with high retention time and having a holding tank with low cost. With that in mind, we erred on the side of lowering capital cost and went with a retention time of one hour. The volume could be found by multiplying the retention time with the feed flow rate (Stream 1). We assumed that a good height to width ratio for a large water tank would be 4:1.

A mass balance was done around the dilution tank to find what amount of seawater would be necessary to dilute the brine to 45,000 ppm salinity (Appendix 3; Table 3). With this the feed flow rate into the tank could be found. It was found that it takes about 20 minutes of retention time for good dilution of the brine in seawater.8 We again assumed that a good height to width ratio for a large water tank would be 4:1.

We assumed a corrosion allowance of 4 mm wall thickness. We calculated wall thickness using vessel pressure, hoop stress, longitudinal stress, and weld factor. Each tank is made with SS316 to protect against corrosion.<ref>Towler G, Sinnott R. Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013.</ref>

<center>
'''Table 2.''' Tank sizes.
[[File:Table_2.jpg|center|600px]]
</center>

=Process Optimization=
<center> '''Table 3.'''Optimization Criteria
[[File:Table_3.jpg|center|400px]]
</center>
In order to optimize the reverse osmosis system, numerous simulations were run. Three representative cases are reported as shown above (Table 3). The objective of the optimization was to minimize the required water price to break even in 10 years. In order to achieve this goal, the balance between the capital cost from the number of the filter elements and utility cost from energy consumption was weighed. Before the energy recovery system was implemented, the general trend was that if we have more filter elements in series, overall recovery increases but recovery rate per filter decreases, and we need more filter elements. While increasing the elements in series increase the required number of filters, and thereby increases the capital cost, we require less energy to produce the same amount of water. After the energy recovery system is implemented, the correlation between the number of elements in series and the required energy becomes more complicated because high pressure concentrate is used to pressurize the feed. However, the trade-off between the utility cost and capital cost still exists. Of the three cases shown above, single stage, five elements RO system turned out to be the most optimal solution because single stage, single element system required higher utility cost and two stage system slightly did not meet the LA standard. Detailed information which was used for the optimization is available in Appendix 3 and in Appendix 6.

=Ethical, Environmental, and Societal Considerations=

There are a few key considerations we have followed to meet ethical, environmental, and societal guidelines and requirements: 
* Meet L.A. Drinking Water Standards
** Correct ion concentration
* Moved near water treatment plant
** Pipes will not corrode from low pH
* Working below pressure limits
* Water will be locally served and meet taste standards
** Via Post Treatment mineralization
* Follows EPA Environmental Salinity Requirements
* Reduced Energy Requirements
** PX Pressure Exchanger
** Optimized RO System
** Reverse Osmosis is itself Energy Saving

For further environmental considerations, see effluent summary in Appendix 4.

=Economic Analysis=

The pricing of the majority of the process was completed using the Aspen Economic Analyzer, however we had determine costs for the ultrafiltration, reverse osmosis membranes, and the ultraviolet treatment was completed separately. This led to a total ISBL of $56 MM, OSBL of $22.4 MM, Engineering Costs of $22.4 MM and a Contingency of $7.8 MM. We also chose a working Capital of $10 MM. This working capital is sufficient because we have no raw materials costs. We also had FCOP of $5.2 MM and VCOP of $3.5 MM.

We originally were planning on pricing our water at $700 per acre-foot, but after doing economic analysis we had a 10 year NPV of $-76.2 MM. After determining did a sensitivity analysis and determined that if our water was priced at $2104 per acre-foot we could have an NPV of $0 after 10 years and if we were able to get the same price of $2257 per acre-foot as the plant in Carlsbad we would have a 10 year NPV of $8.3 MM which increases to $41.1 MM at 20 years. For more detailed data about the economic analysis see Appendix 6.

=Conclusion and Recommendations=

Our mission was to create an economically viable process that will desalinate water for human consumption in the L.A. County. We have found after research and optimization that the best process to undergo desalination is a continuous single stage, five elements reverse osmosis system. Our process can meet the required L.A. drinking water standard and we have set up plans for post treatment at a local water reclamation site. While meeting the drinking water standards, we also met our ethical, environmental, and social standards.

Nevertheless, at a sale price of $700 per acre-foot, our proposed desalination plant in Los Angeles, California looks to be a very expensive project to undergo. Without any sort of government assistance to carry the heavy capital costs of such a project, we would suggest the plant not be built. However, after performing a sensitivity analysis at higher (and less conservative) price points we reached a breakeven point in 10 years with a sale price of $2104 per acre-foot of drinking water. Research has shown that the droughts striking California look only to get worse in the near future. With limited freshwater resources, there is no option other than to consider desalination. Therefore it is reasonable to expect we could get near the same price deal as local desalination plants like the Carlsbad Plant ($2257 per acre-foot) As we have met all design constraints for this project, we hope to convince policymakers to adopt our project at least $2104 per acre-foot water. If this deal is reached, we recommend that this project and design be funded.

However, there are ways to improve our design with future work. Right now all the pipes diameters are based on having a flowrate of 1.5 m/s. But in reality, we’d need to use standard piping sizes to get more accurate pricing and energy costs. Also, in reality we would be purchasing some of these vessels in bulk quantities (i.e. bulk filters) so the price could be lowered with negotiation and economies of scale. We also have pumps that are providing very little pressure for their high cost. We can remove these pumps and simply run at higher pressures earlier. We can potentially reduce early energy costs by building a “floating” desalination plant in the sea to reduce pumping distance from the well.

=Appendices=

==Appendix 1: Design Basis==
[[File:Appendix_1.png|thumb|border|center|800px|middle|]]

==Appendix 2: Process Flow Diagram==
[[File:Appendix_2.png|center]]
==Appendix 3: Mass and Energy Balance Tables==
'''Table 1'''. Mass balance on 1 stage 5 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T1.png|center]]

'''Table 2'''. Mass balance on 1 stage 1 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T2.png|center]]

'''Table 3'''. Mass balance on 2 stage 5,5 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T3.png|center]]

'''Table 4'''. 1 stage 5 elements CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T4.png|center]]

'''Table 5'''. 1 stage 1 element CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T5.png|center]]

'''Table 6'''. 2 stages 5,5 elements CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T6.png|center]]

'''Table 7'''. Energy balance on 1 stage 5 elements RO system
[[File:Appendix_3_T7.png|center]]

==Appendix 4: Effluent Summary==
[[File:Appendix_4.png|center]]

==Appendix 5: Cost of Production==
[[File:Appendix_5.png|center]]

==Appendix 6: Economic Anlaysis==
===Case I ===
Water is $700; RO with 1 stage and 5 elements
[[File:Appendix_6_C1.png|center]]

===Case II===
Water is $1400; RO with 1 stage and 5 elements
[[File:Appendix_6_C2.png|center]]

===Case III: Carlsbad ===
Water is $2257; RO with 1 stage and 5 elements
[[File:Appendix_6_C3.png|center]]

===Case IV: Breakeven (NPV = 0 at 10 years)===
Water is $2104; RO with 1 stage and 5 elements
[[File:Appendix_6_C4.png|center]]

===Case V: Breakeven (NPV = 0 at 10 years)===
Water is $2112; RO with 2 stages and 5,5 elements
[[File:Appendix_6_C5.png|center]]

===Case VI: Breakeven (NPV = 0 at 10 years)===
Water is $2138 with 1 stage and 1 element
[[File:Appendix_6_C6.png|center]]

==Appendix 7: ROSA Simulation Outputs==

'''Table 1'''. System detail of 1 stage 5 elements RO system
[[File:Appendix_7_T1.png|center]]

'''Table 2'''. System detail of 1 stage 1 elements RO system
[[File:Appendix_7_T2.png|center]]

'''Table 3'''. System detail of 2 stage 5,5 elements system
[[File:Appendix_7_T3.png|center]]

'''Table 4'''. 1 stage 5 elements RO system concentrations and pH in each stream
[[File:Appendix_7_T4.png|center]]

'''Table 5'''. 1 stage 1 element RO system concentration and pH in each stream
[[File:Appendix_7_T5.png|center]]

'''Table 6'''. 2 stage 5,5 elements RO system concentrations and pH in each stream
[[File:Appendix_7_T6.png|center]]

=References=
<references/>

Desalination - Team E

2016-03-12T03:29:46Z

Bts511: /* Appendix 1: Design Basis */

Group E Corporation

Authors: Hassan Ali, Woo Soo Choe, Brett Sleyster, Jake Stolley

Instructors: Fengqi You, David Wegerer

March 11, 2016

=Executive Summary=

Outlined in this report is the project model report for the proposed desalination plant in Los Angeles, California which will provide drinking water to local citizens using reverse osmosis process. This location was chosen due to the region’s propensity for droughts in recent years. The fresh water resources in California have plummeted as water prices have risen 3-4% each year. This is only expected to continue. Other companies have worked deals with the Californian government to create desalination plants that provide a new fresh water drinking source. This suggests that there is a strong market here with potential growth.

The plant will have an output of 15 million gallons per day, which is about medium size relative to other plants. Thus the build time will be about one year. The plant will run for 30 years, with 350 days per year of use. Shutdown time will be used for cleaning and maintenance of equipment. The plant will be built on Long Beach, 500 meters from the coast nearby the Long Beach Water Reclamation Plant (LBWP). The input will withdraw 70.5 million gallons per day of seawater that will be used for reverse osmosis and dilution of brine.

Reverse osmosis was chosen as the desalination method for its low energy cost and flexibility. We designed a process that will pretreat the seawater to remove waste, colloids, bacteria, and membrane-damaging scalants. Once this seawater is treated is pumped to heavy pressures through a reverse osmosis system. After doing process optimization under various system parameter changes, we found a continuous, single stage, five element reverse osmosis process to be best. It produces a near 30% recovery. We reduce energy costs by including a pressure exchanger between our concentrate outlet and adjusted feed into the reverse osmosis system.

After performing mass and energy balances, we were able to size equipment and find our permeate composition. It turns out our process will meet L.A. standards for drinking water quality. However, post treatment will be done by the local LBWP to make the water taste better and be non-caustic to pipes. This will reduce “in-house” costs. Many environmental considerations were made, to meet our environmental standards. We meet EPA salinity requirements for our brine waste by diluting it. We also don’t use a lot of energy and thus reduce our fossil fuel consumption and CO2 generation. To be safe, we work well under max pressure limits for our equipment. We make sure we are ethical and meet all standards that our local county requires of us.

After an economic evaluation at a conservative well-water price of $700 per acre-foot, we realize that our revenue is way too low to cover our costs. Capital costs are especially large with regards to our large water tanks. We find that we reach a 10 year NPV breakpoint of $2104 per acre-foot after performing a sensitivity analysis. The Carlsbad Plant in San Diego worked out a deal for $2257 per acre-foot.

As we have met all design constraints for this project, we hope to convince policymakers to adopt our project at least $2104 per acre-foot water. If this deal is reached, we recommend that this project and design be funded. There is future work to be done to improve costs such as performing pump optimization. However, this work is minimal and won’t harm our bottom line if we reach our proposed price points.

__TOC__

=Introduction=

Due to a combination of various factors--including rising human population, climate change, increased energy consumption, and land erosion--the availability of fresh water as a resource for human consumption has been dwindling. This has had a negative impact on the world, but this impact has been especially felt in areas with natural shortages in fresh water. Desert regions, such as Saudi Arabia, have resorted to desalinating seawater at the coast and transporting up to their cities. This encompasses their entire freshwater system as very few natural water sources remain. In the United States there is a region which has shown increasing signs of following in the footsteps of Saudi Arabia and becoming fully reliant on desalination: Los Angeles, California.

=Location and Market Analysis=

In this project, Los Angeles was chosen as the location our plant, because the area has been subject to a severe drought for the last four years, as shown in Figure 1. The rapid decrease in the available water has caused permanent damages to the ground, decreased the amount of water in the reservoir from 45% to 25%, and reduced the amount of hydroelectric power produced in the state. Such deficiency in usable water has increased the water prices in Los Angeles by 6-7% annually for the last decade and is expected to rise by 3.4% annually for typical customers for the next five years. For customers who use more than 20000 gallons a year, the bill is expected to increase by 34% by 2021. This price increase is occurring because of both drought and maintenance. <ref>Pacific Institute: The California Drought. 2014.</ref>

[[File:Figure_1.jpg|right|200px|frame|Figure 1. California Drought Index.]]
As of March 2015 the Metropolitan Water District of Southern California was willing to pay as much as $700 per acre-foot of water to farmers in the Sacramento Valley, so we assume this will be the target price of the water produced by the this desalination plant.<ref>Hessel, Phil. “Dry Southern California Offers Northern Farmers Top Dollar for Water”. NBC News. March 18, 2015. </ref>

Currently, there is a plant in Carlsbad which supplies drinking water to the San Diego area. The plant can produce up to 50 million gallons per day and is the nation's largest water desalination plant. The water from this plant costs $2257 per acre-foot for the first 48000 acre-feet and about $2000 for any additional acre-foot. The county signed a contract with this plant to purchase the first 48000 acre-feet of water for the next 30 years.<ref>Carlsbad Desalination Plant. “Project Overview”. January 2016.</ref> Also, there is another desalination plant in Tampa that produces about 25 million gallons of water per day, and over 2000 other public and industrial plants with production capacity of over 300000 gallons per day. When conducting the economic analysis our plant, such numbers presented by other plants must be considered, so our project can yield positive profit.<ref>Leven, Rachel. “U.S. Desalination Industry Grows Since 2000; See as Essential to Meeting Supply Needs”. Bloomberg BNA. August 21, 2013.</ref>

The profitability of the project is naturally heavily dependent on the price of water in California. If the drought ends, water supply will increase and the water price will drop significantly, but if the drought continues, the price of water in California could rise even more rapidly and allow desalination plants to be more profitable.<ref>Southern California Public Radio. “Where is California water use decreasing?”. December 2015.</ref> It may be difficult to make this plant profitable without a guaranteed purchase or other subsidy from the local governments because the presence of pre-existing plants maintains the water price relatively low. However, external sources of profit such as guaranteed purchase or subsidy are expected because the ramifications of not having desalination plant or other sources of water may be significantly more expensive than the subsidizing a desalination plant assuming the drought continues.<ref>Weiser, Matt. “Could desalination solve California’s water problem?”. The Sacramento Bee. October 18, 2014.</ref>

=Design Basis and Technical Approach=

[[File:Figure_2.jpg|left|200px|frame|Figure 2. Vertical Well Feed]]
Considering the number of pre-existing desalination plant and the potential risk of the drought ending, our plant will be tentatively designed to produce 15 million gallons of drinkable water per day to homes and businesses in the Los Angeles county area. With a recovery rate of near 30%, and a requirement for diluting any brine, the input feed flowrate will be 70.5 million gallons per day. Following the similarly built Carlsbad plant, we expect the plant to last 30 years. The project building timeline is one year based on our relatively small plant size. We will run for 350 days a year while shutting down for cleaning and maintenance once every 3 months (based on DOW standards).<ref name = "Bates">Bates, Wayne. “Cleaning your RO”. Hydranautics. </ref> The plant will be built 500 meters from the coastline on Long Beach. A borehole will be drilled 10 meters into the ground. At this depth there should be seawater aquifers below the coast that regularly recharges itself with seawater from the ocean. A vertical well containing intake holes will be put into the borehole. The intake holes will be equipped with large size filters that can easily be replaced. At the surface, the well will be connected to a large-scale pump in a well pump house. From there, the seawater will be horizontally pumped into a seawater tank in the desalination plant. This will be considered the start of our process. The process ends after pretreatment, desalination, post treatment, and brine treatment are completed. Appendix 1 shows more details on the Design Basis.

The feed composition is assumed to be the standard seawater “average” based on the DOW Chemical ROSA 9.0 Technical Manual in Table 1. However, in reality the seawater composition at the plant site will have to be chemically tested. The final product will be based on L.A. standards for drinking water quality (Table 1). The specifications in the table are upper limits that we cannot exceed. We will aim to reach the salt level standards but any remineralization and pH treatment will be exported and paid for due to the added complexity and cost in making such a system in-house.

Due to the acidity of the desalinated water coming out of the reverse osmosis (RO) system (pH ~ 6), the pumps and pipes carrying the desalinated water for post treatment will likely corrode over time. We have decided to place the desalination plant by the Long Beach Water Reclamation Plant (about 50 meters away) to prevent pipe corrosion.

[[File:Table_1.jpg|center|200px|frame|Table 1. Concentration of the standard seawater, the output from the RO system, and LA standard.]]

=Desalination Overview and Alternatives=

We first began our research into the current landscape of the desalination industry, particularly of plants that were built in the United States. We discovered there were many processes but the main ones included: 
* Membrane Separation (Forward Osmosis and Reverse Osmosis)
* Thermal Separation (Multi Stage Flash Distillation)
* Phase Change Separation (Freeze Separation)
Each process was analyzed for their pros and cons. For our project design, we decided to go with reverse osmosis as our desalination process. This is because it is the most commonly used process in the industry, and for good reason. Unlike thermal separation, which requires huge amounts of energy to supply heat, reverse osmosis is very energy efficient. This reduces energy costs that would be spread across the plant’s lifetime.<ref name = "Republic">Republic of Mauritius NY. "Technical Aspects of Desalination Plant". Sustainable Sanitation and Water Management. </ref>

Forward osmosis is a newer technology being developed that also uses membranes to separate water and salt and may be more energy efficient overall than reverse osmosis. This is because forward osmosis doesn’t require a pump to pressurize seawater through the membranes, but rather leaves a higher salt concentrated solution (AKA draw solution) on the other side of the membrane for the seawater to be drawn into. Then a second step is required to regenerate the concentrated draw solution and produce purified water.<ref>Mallinson, Alissa. "Study Shows Forward Osmosis Desalination Not Energy Efficient." MIT News [Boston]. July 23, 2014.</ref> We decided to move away from forward osmosis because it is a process that is still being refined and developed for desalination and lacks the data, programs, and technologies that reverse osmosis users benefit from. Additionally, a recent study from MIT has claimed that due to the inherent nature of forward osmosis requiring that seawater be drawn to a higher concentrated salt solution before the water and salt is separated, it cannot possibly be more energy efficient than reverse osmosis.<ref>O'Hern, Sean C., Karnik, Rohit. Nanofiltration across Defect-Sealed Nanoporous Monolayer Graphene. American Chemical Society 15.5 (2014); 3254-260. </ref>

Other processes such as freeze separation were too restrictive in their feed and process demands and required a lot of capital. Ultimately, reverse osmosis is simply the best process that exists for desalination and has been ranked the best when performance is evaluated in terms of technical, environmental, and economic aspects. It is a process that is flexible in water quantity, operation start-up and shut-off, and site location.

Nevertheless there are disadvantages to reverse osmosis that must be considered. The feed seawater must not be too salty and too low quality. The pressure to pump the seawater is enormous. The construction time is long for large plants. The capital costs are also high when compared with other processes such as thermal separation. But all of these disadvantages are ultimately outweighed by the advantages reverse osmosis provides.<ref name = "Republic"/>

=Process Overview and Alternatives=

See Appendix 2 (Figure 1) for the Process Flow Diagram (PFD). The process itself is detailed below. The stream flow rates are in Appendix 2 (Table 1). We assumed 10 meters of pipe for every stream except the inlet and outlet streams of the entire PFD.

==Pretreatment==

After the 10 meter vertical well pumps seawater from the Long Beach coast, the feed is pumped 500 meters to a seawater tank that makes it easy to shut-down the system and reduce the large volume of flow. The level control keeps the seawater from overfilling the tank. From there, part of the seawater goes to brine treatment while the rest go to pretreatment. The purpose of pretreatment in the desalination process is to prevent equipment damage by removing particles and microorganisms from the water. A variety of pollutants -- including calcium carbonate, sulfates, silica, clay, bacteria, and other colloids and microorganisms -- can disrupt the reverse osmosis process. Media filtration uses fine grains such as sand or anthracite to filter the particles. We used a large filter to catch larger particle material (e.g. sand). However, simple filtration is not enough to prevent colloid fouling. An ultrafilter (i.e. microfiltration) is used to remove virtually all suspended matter, effectively preventing colloid fouling. The costs associated with ultrafiltration are not prohibitive in terms of the project’s budget.<ref name = "Desalination">Desalination Pre-treatment. Lenntech.</ref>

Reverse osmosis systems can also be disrupted by the growth of microbiological films of bacteria on the film and pipe walls. They are typically treated as colloids during pretreatment, and most organisms are removed from the water by ultrafiltration, but this does not remove the smallest of organic particles. DBNPA is the most effective biocide, but these aren’t often used and have to be removed from the water. Chlorination is an effective way to remove bacteria, but it can cause degradation of the membrane, and must be removed before the reverse osmosis occurs. In our process we will use UV irradiation to remove any organic material that passes through the microfilters.<ref name = "Desalination"/>

Scaling can reduce the flux across the membrane and can build up in pipe corners, causing backup. It is caused by inorganic salts -- specifically sulfates and calcium and magnesium carbonate -- present in the water. Scaling can be prevented and/or controlled with acids, antiscalants, or ion exchange softening. Antiscalants are good for removing sulfates, but not much else. Acidification will only be effective at removing CaCO3 from water but, the decreased pH may deteriorate the membrane and pipes, so it may require future research.<ref name = "Desalination"/> Ion exchange is very expensive and requires more capital to undergo, but is overall the more effective option.

The ion exchanger is a pretreatment system which acts as a method to fight membrane scaling in reverse osmosis. The ion exchanger works by using sodium carboxylate as a resin which will react with Chlorine and Magnesium ions, removing them from the feed. The sodium is substituted out and enters the feed stream, increasing the overall sodium level and TDS. The lack of magnesium and chlorine ions means that there is nothing to react with the carbonates to form scaling salts on the membranes.

==Desalination (Reverse Osmosis)==

Two pumps (P-105 & P-106) take the adjusted feed seawater up to 58 bar. However, before this pumping occurs, about 70% of the original stream gets separated and heads into a PX Pressure Exchanger. This separated stream later joins together with the pumped stream to enter the reverse osmosis system at the required pressure of 58 bar. The streams can combine together as they have equal pressure. It is the concentrate stream from the output of the reverse osmosis system that transfers its energy to the adjusted feed in the PX Pressure Exchanger.

[[File:Figure_3.jpg|left|300px|frame|Figure 3. Inside look of pressure exchanger]]
The pressure of the concentrate stream is about 57 bar and it is fed into an exchanger which works by using vessel that resembles a small pipe within a rotating wheel. The low pressure (LP) feed enters the vessel at 1 bar (see Figure 3). The feed turns and is pushed out into an outlet pipe by the high pressure (HP) concentrate stream at ~57 bar. The feed is now a HP stream via this exchanged energy. The now LP concentrate stream turns and is pushed out into an outlet pipe by the LP feed entering the vessel.

98% of the energy within the concentrate stream can be recovered via the PX Pressure Exchanger.<ref>Desalination Product and Applications Catalog 2013. “Advance Technologies for Seawater Reverse Osmosis and Brackish Water Applications”. Energy Recovery, Inc. 2012-2014.</ref >Thus, the now HP feed stream is not quite 58 bar and will need to be pumped by a small booster pump (P-107).

When designing a reverse osmosis system, we need to make decisions on what specific kind of reverse osmosis system will be implemented. When making the decision, we investigated the advantages of different types of reverse osmosis among the following options: batch vs. continuous process, number of stages, and number of elements.<ref>FILMTEC Reverse Osmosis Membranes. “Water & Process Solutions Technical Manual.” Dow Chemical Company.</ref> Our decision was made based on process optimization which is described later in this report. The end result of that optimization was to go with a single stage, continuous process with five elements in series.

=Post Treatment=

The permeate is sent to post treatment after the salt content is analyzed and controlled if too high or too low due to any unexpected errors. As stated above, the post treatment process will be done 50 meters from our plant by the Long Beach Water Reclamation Plant. The reason we aren’t performing post treatment “in-house” is due to the increased cost it would take for us to build and run the equipment for post treatment. We have the option to use the government’s water treatment facilities that already exist and our 15 million gallons per day output is a relatively small addition to the large Long Beach Water Reclamation Plant. We will assume any costs to post treatment will be pushed to the consumer in the price of purchasing the drinking water from the Long Beach Water Reclamation Plant.

Post treatment at any facility works to bring water to “taste standards”. The water that is the permeate of our process is already meeting L.A. drinking standards as shown above in Table 1 after our mass balance calculations around the reverse osmosis system (Appendix 3, Table 1). The post treatment increases the pH from 6.3 to 7.3 with the addition of base and then adds minerals like fluoride, calcium, and magnesium back in to improve taste.

==Brine Treatment and Byproduct Waste==

A natural consequence of a reverse osmosis procedure is that a lot of salt will be leftover from the membrane separation in a very concentrated form. This is called “brine” and will have to be taken care of. According to EPA requirements, the salt concentration of water entering the environment (via land or sea) cannot be higher than 40,000 ppm.8 The salt concentration of our brine is much higher (around 70,000 ppm). Thus, the brine must be removed or treated. We decided to treat the brine by diluting it as removing brine can be difficult and may require lots of land (e.g. evaporation ponds).

After the concentrate exchanges energy in the energy exchanger, it becomes a low pressure stream. It is pumped into a dilution tank and mixes with 18.2 million gallons per day of seawater from the inlet to become diluted brine which is pumped 600 meters out to sea.

The waste byproducts are the solid buildup on the filters and membranes. The filters and membranes also have a set lifetime and thus will also need to be removed. The waste is all non-hazardous and thus can be put in the municipal dump.

=Process Details and Equipment Sizing=

==Filters and UV Treatment [Pretreatment]==
The cost of our ultrafiltration system is estimated using data from the University of Tennessee at Knoxville.<ref>Qiang He and R. Bruce Robinson Membrane Filtration Systems. “Membrane Module Costs.” University of Tennessee at Knoxville. December 2014.</ref> It estimated the cost of the cost of the ultrafiltration $0.057 per 1000 gallons and the construction and capital costs are $0.39 per gallon per day. The cost of UV treatment is based off of units available from buyultraviolet.com for units that can treat about 450 gallons per minute.<ref>Atlantic Ultraviolet Corporation. “Megatron Water Disinfection System.”</ref> Due to the water having to be in relatively small pipes for treatment we do not expect the costs to scale very much, so we believe these units to be an accurate cost estimate.

==Ion Exchanger [Pretreatment] ==

The ion exchanger was sized using data from DOW. Our process will use 60 vessels with ion exchange beds of volume 5.73 m3. With a flow rate of 30 bed volumes/hour (10,312.2 m3/hr) the system of beds has a capacity of 125% of the feed. This allows for periodic regeneration of the resin, a weak acid cation resin (DOW MAC-3 Resin). This weak cation resin will remove cationic calcium and magnesium from the water to prevent scaling. The mass flow was calculated via a mass balance on the ion exchanger (Appendix 3).

==Reverse Osmosis System [Desalination]==

Each of the 2607 Reverse Osmosis SW30HRLE-440i [Seawater 30 High Rejection Low Energy 440 ft2 interlocking cap membrane] units have the same dimensions. A, the net length, is 1016 mm. B, the total length including the endcap, is 1029 mm. C, the diameter of the filter element, is 201mm. D, the diameter of the permeate pathway, is 29 mm. For a single stage, we have five of the elements in series so the overall length of a single stage is expected to be roughly 5145 mm. The membrane material is polyamide thin film composite. It will last three years.
Operating Conditions [ROSA 9.0]: 
* Temperature: 30oC 
* Pressure: 58 bar 
* Flow Rate: 52.3 Million gal per day 
* Recovery: 28.7%

[[File:Figure_4.png|center|400px|frame|Figure 4. Diagram of Dow FILMTEC SW30HRLE-440i]]
Cleaning Schedule: 
* Conservative Estimate: Once every three months (three days available for cleaning)<ref name = "Bates"/> 
* DOW FILMTEC Cleaning Conditions 
** The normalized permeate flow drops 10%
* The normalized salt passage increases 5 - 10%

==Seawater Holding Tank and Dilution Tank [Brine Treatment]==

The initial seawater holding tank works as a buffer to prevent surges and help perform shutdowns quickly and easily. We expect to run continuously for 350 days a year with the rest of the available time spread out and used for shutting down the plant for cleaning membranes, filters, and the fixing of unexpected errors. There is a big tradeoff between having a holding tank with high retention time and having a holding tank with low cost. With that in mind, we erred on the side of lowering capital cost and went with a retention time of one hour. The volume could be found by multiplying the retention time with the feed flow rate (Stream 1). We assumed that a good height to width ratio for a large water tank would be 4:1.

A mass balance was done around the dilution tank to find what amount of seawater would be necessary to dilute the brine to 45,000 ppm salinity (Appendix 3; Table 3). With this the feed flow rate into the tank could be found. It was found that it takes about 20 minutes of retention time for good dilution of the brine in seawater.8 We again assumed that a good height to width ratio for a large water tank would be 4:1.

We assumed a corrosion allowance of 4 mm wall thickness. We calculated wall thickness using vessel pressure, hoop stress, longitudinal stress, and weld factor. Each tank is made with SS316 to protect against corrosion.<ref>Towler G, Sinnott R. Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013.</ref>

<center>
'''Table 2.''' Tank sizes.
[[File:Table_2.jpg|center|600px]]
</center>

=Process Optimization=
<center> '''Table 3.'''Optimization Criteria
[[File:Table_3.jpg|center|400px]]
</center>
In order to optimize the reverse osmosis system, numerous simulations were run. Three representative cases are reported as shown above (Table 3). The objective of the optimization was to minimize the required water price to break even in 10 years. In order to achieve this goal, the balance between the capital cost from the number of the filter elements and utility cost from energy consumption was weighed. Before the energy recovery system was implemented, the general trend was that if we have more filter elements in series, overall recovery increases but recovery rate per filter decreases, and we need more filter elements. While increasing the elements in series increase the required number of filters, and thereby increases the capital cost, we require less energy to produce the same amount of water. After the energy recovery system is implemented, the correlation between the number of elements in series and the required energy becomes more complicated because high pressure concentrate is used to pressurize the feed. However, the trade-off between the utility cost and capital cost still exists. Of the three cases shown above, single stage, five elements RO system turned out to be the most optimal solution because single stage, single element system required higher utility cost and two stage system slightly did not meet the LA standard. Detailed information which was used for the optimization is available in Appendix 3 and in Appendix 6.

=Ethical, Environmental, and Societal Considerations=

There are a few key considerations we have followed to meet ethical, environmental, and societal guidelines and requirements: 
* Meet L.A. Drinking Water Standards
** Correct ion concentration
* Moved near water treatment plant
** Pipes will not corrode from low pH
* Working below pressure limits
* Water will be locally served and meet taste standards
** Via Post Treatment mineralization
* Follows EPA Environmental Salinity Requirements
* Reduced Energy Requirements
** PX Pressure Exchanger
** Optimized RO System
** Reverse Osmosis is itself Energy Saving

For further environmental considerations, see effluent summary in Appendix 4.

=Economic Analysis=

The pricing of the majority of the process was completed using the Aspen Economic Analyzer, however we had determine costs for the ultrafiltration, reverse osmosis membranes, and the ultraviolet treatment was completed separately. This led to a total ISBL of $56 MM, OSBL of $22.4 MM, Engineering Costs of $22.4 MM and a Contingency of $7.8 MM. We also chose a working Capital of $10 MM. This working capital is sufficient because we have no raw materials costs. We also had FCOP of $5.2 MM and VCOP of $3.5 MM.

We originally were planning on pricing our water at $700 per acre-foot, but after doing economic analysis we had a 10 year NPV of $-76.2 MM. After determining did a sensitivity analysis and determined that if our water was priced at $2104 per acre-foot we could have an NPV of $0 after 10 years and if we were able to get the same price of $2257 per acre-foot as the plant in Carlsbad we would have a 10 year NPV of $8.3 MM which increases to $41.1 MM at 20 years. For more detailed data about the economic analysis see Appendix 6.

=Conclusion and Recommendations=

Our mission was to create an economically viable process that will desalinate water for human consumption in the L.A. County. We have found after research and optimization that the best process to undergo desalination is a continuous single stage, five elements reverse osmosis system. Our process can meet the required L.A. drinking water standard and we have set up plans for post treatment at a local water reclamation site. While meeting the drinking water standards, we also met our ethical, environmental, and social standards.

Nevertheless, at a sale price of $700 per acre-foot, our proposed desalination plant in Los Angeles, California looks to be a very expensive project to undergo. Without any sort of government assistance to carry the heavy capital costs of such a project, we would suggest the plant not be built. However, after performing a sensitivity analysis at higher (and less conservative) price points we reached a breakeven point in 10 years with a sale price of $2104 per acre-foot of drinking water. Research has shown that the droughts striking California look only to get worse in the near future. With limited freshwater resources, there is no option other than to consider desalination. Therefore it is reasonable to expect we could get near the same price deal as local desalination plants like the Carlsbad Plant ($2257 per acre-foot) As we have met all design constraints for this project, we hope to convince policymakers to adopt our project at least $2104 per acre-foot water. If this deal is reached, we recommend that this project and design be funded.

However, there are ways to improve our design with future work. Right now all the pipes diameters are based on having a flowrate of 1.5 m/s. But in reality, we’d need to use standard piping sizes to get more accurate pricing and energy costs. Also, in reality we would be purchasing some of these vessels in bulk quantities (i.e. bulk filters) so the price could be lowered with negotiation and economies of scale. We also have pumps that are providing very little pressure for their high cost. We can remove these pumps and simply run at higher pressures earlier. We can potentially reduce early energy costs by building a “floating” desalination plant in the sea to reduce pumping distance from the well.

=Appendices=

==Appendix 1: Design Basis==
[[File:Appendix_1.png|thumb|border|center|600px|middle|]]

==Appendix 2: Process Flow Diagram==
[[File:Appendix_2.png|center]]
==Appendix 3: Mass and Energy Balance Tables==
'''Table 1'''. Mass balance on 1 stage 5 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T1.png|center]]

'''Table 2'''. Mass balance on 1 stage 1 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T2.png|center]]

'''Table 3'''. Mass balance on 2 stage 5,5 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T3.png|center]]

'''Table 4'''. 1 stage 5 elements CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T4.png|center]]

'''Table 5'''. 1 stage 1 element CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T5.png|center]]

'''Table 6'''. 2 stages 5,5 elements CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T6.png|center]]

'''Table 7'''. Energy balance on 1 stage 5 elements RO system
[[File:Appendix_3_T7.png|center]]

==Appendix 4: Effluent Summary==
[[File:Appendix_4.png|center]]

==Appendix 5: Cost of Production==
[[File:Appendix_5.png|center]]

==Appendix 6: Economic Anlaysis==
===Case I ===
Water is $700; RO with 1 stage and 5 elements
[[File:Appendix_6_C1.png|center]]

===Case II===
Water is $1400; RO with 1 stage and 5 elements
[[File:Appendix_6_C2.png|center]]

===Case III: Carlsbad ===
Water is $2257; RO with 1 stage and 5 elements
[[File:Appendix_6_C3.png|center]]

===Case IV: Breakeven (NPV = 0 at 10 years)===
Water is $2104; RO with 1 stage and 5 elements
[[File:Appendix_6_C4.png|center]]

===Case V: Breakeven (NPV = 0 at 10 years)===
Water is $2112; RO with 2 stages and 5,5 elements
[[File:Appendix_6_C5.png|center]]

===Case VI: Breakeven (NPV = 0 at 10 years)===
Water is $2138 with 1 stage and 1 element
[[File:Appendix_6_C6.png|center]]

==Appendix 7: ROSA Simulation Outputs==

'''Table 1'''. System detail of 1 stage 5 elements RO system
[[File:Appendix_7_T1.png|center]]

'''Table 2'''. System detail of 1 stage 1 elements RO system
[[File:Appendix_7_T2.png|center]]

'''Table 3'''. System detail of 2 stage 5,5 elements system
[[File:Appendix_7_T3.png|center]]

'''Table 4'''. 1 stage 5 elements RO system concentrations and pH in each stream
[[File:Appendix_7_T4.png|center]]

'''Table 5'''. 1 stage 1 element RO system concentration and pH in each stream
[[File:Appendix_7_T5.png|center]]

'''Table 6'''. 2 stage 5,5 elements RO system concentrations and pH in each stream
[[File:Appendix_7_T6.png|center]]

=References=
<references/>

Desalination - Team E

2016-03-12T03:29:12Z

Bts511: /* Appendix 1: Design Basis */

Group E Corporation

Authors: Hassan Ali, Woo Soo Choe, Brett Sleyster, Jake Stolley

Instructors: Fengqi You, David Wegerer

March 11, 2016

=Executive Summary=

Outlined in this report is the project model report for the proposed desalination plant in Los Angeles, California which will provide drinking water to local citizens using reverse osmosis process. This location was chosen due to the region’s propensity for droughts in recent years. The fresh water resources in California have plummeted as water prices have risen 3-4% each year. This is only expected to continue. Other companies have worked deals with the Californian government to create desalination plants that provide a new fresh water drinking source. This suggests that there is a strong market here with potential growth.

The plant will have an output of 15 million gallons per day, which is about medium size relative to other plants. Thus the build time will be about one year. The plant will run for 30 years, with 350 days per year of use. Shutdown time will be used for cleaning and maintenance of equipment. The plant will be built on Long Beach, 500 meters from the coast nearby the Long Beach Water Reclamation Plant (LBWP). The input will withdraw 70.5 million gallons per day of seawater that will be used for reverse osmosis and dilution of brine.

Reverse osmosis was chosen as the desalination method for its low energy cost and flexibility. We designed a process that will pretreat the seawater to remove waste, colloids, bacteria, and membrane-damaging scalants. Once this seawater is treated is pumped to heavy pressures through a reverse osmosis system. After doing process optimization under various system parameter changes, we found a continuous, single stage, five element reverse osmosis process to be best. It produces a near 30% recovery. We reduce energy costs by including a pressure exchanger between our concentrate outlet and adjusted feed into the reverse osmosis system.

After performing mass and energy balances, we were able to size equipment and find our permeate composition. It turns out our process will meet L.A. standards for drinking water quality. However, post treatment will be done by the local LBWP to make the water taste better and be non-caustic to pipes. This will reduce “in-house” costs. Many environmental considerations were made, to meet our environmental standards. We meet EPA salinity requirements for our brine waste by diluting it. We also don’t use a lot of energy and thus reduce our fossil fuel consumption and CO2 generation. To be safe, we work well under max pressure limits for our equipment. We make sure we are ethical and meet all standards that our local county requires of us.

After an economic evaluation at a conservative well-water price of $700 per acre-foot, we realize that our revenue is way too low to cover our costs. Capital costs are especially large with regards to our large water tanks. We find that we reach a 10 year NPV breakpoint of $2104 per acre-foot after performing a sensitivity analysis. The Carlsbad Plant in San Diego worked out a deal for $2257 per acre-foot.

As we have met all design constraints for this project, we hope to convince policymakers to adopt our project at least $2104 per acre-foot water. If this deal is reached, we recommend that this project and design be funded. There is future work to be done to improve costs such as performing pump optimization. However, this work is minimal and won’t harm our bottom line if we reach our proposed price points.

__TOC__

=Introduction=

Due to a combination of various factors--including rising human population, climate change, increased energy consumption, and land erosion--the availability of fresh water as a resource for human consumption has been dwindling. This has had a negative impact on the world, but this impact has been especially felt in areas with natural shortages in fresh water. Desert regions, such as Saudi Arabia, have resorted to desalinating seawater at the coast and transporting up to their cities. This encompasses their entire freshwater system as very few natural water sources remain. In the United States there is a region which has shown increasing signs of following in the footsteps of Saudi Arabia and becoming fully reliant on desalination: Los Angeles, California.

=Location and Market Analysis=

In this project, Los Angeles was chosen as the location our plant, because the area has been subject to a severe drought for the last four years, as shown in Figure 1. The rapid decrease in the available water has caused permanent damages to the ground, decreased the amount of water in the reservoir from 45% to 25%, and reduced the amount of hydroelectric power produced in the state. Such deficiency in usable water has increased the water prices in Los Angeles by 6-7% annually for the last decade and is expected to rise by 3.4% annually for typical customers for the next five years. For customers who use more than 20000 gallons a year, the bill is expected to increase by 34% by 2021. This price increase is occurring because of both drought and maintenance. <ref>Pacific Institute: The California Drought. 2014.</ref>

[[File:Figure_1.jpg|right|200px|frame|Figure 1. California Drought Index.]]
As of March 2015 the Metropolitan Water District of Southern California was willing to pay as much as $700 per acre-foot of water to farmers in the Sacramento Valley, so we assume this will be the target price of the water produced by the this desalination plant.<ref>Hessel, Phil. “Dry Southern California Offers Northern Farmers Top Dollar for Water”. NBC News. March 18, 2015. </ref>

Currently, there is a plant in Carlsbad which supplies drinking water to the San Diego area. The plant can produce up to 50 million gallons per day and is the nation's largest water desalination plant. The water from this plant costs $2257 per acre-foot for the first 48000 acre-feet and about $2000 for any additional acre-foot. The county signed a contract with this plant to purchase the first 48000 acre-feet of water for the next 30 years.<ref>Carlsbad Desalination Plant. “Project Overview”. January 2016.</ref> Also, there is another desalination plant in Tampa that produces about 25 million gallons of water per day, and over 2000 other public and industrial plants with production capacity of over 300000 gallons per day. When conducting the economic analysis our plant, such numbers presented by other plants must be considered, so our project can yield positive profit.<ref>Leven, Rachel. “U.S. Desalination Industry Grows Since 2000; See as Essential to Meeting Supply Needs”. Bloomberg BNA. August 21, 2013.</ref>

The profitability of the project is naturally heavily dependent on the price of water in California. If the drought ends, water supply will increase and the water price will drop significantly, but if the drought continues, the price of water in California could rise even more rapidly and allow desalination plants to be more profitable.<ref>Southern California Public Radio. “Where is California water use decreasing?”. December 2015.</ref> It may be difficult to make this plant profitable without a guaranteed purchase or other subsidy from the local governments because the presence of pre-existing plants maintains the water price relatively low. However, external sources of profit such as guaranteed purchase or subsidy are expected because the ramifications of not having desalination plant or other sources of water may be significantly more expensive than the subsidizing a desalination plant assuming the drought continues.<ref>Weiser, Matt. “Could desalination solve California’s water problem?”. The Sacramento Bee. October 18, 2014.</ref>

=Design Basis and Technical Approach=

[[File:Figure_2.jpg|left|200px|frame|Figure 2. Vertical Well Feed]]
Considering the number of pre-existing desalination plant and the potential risk of the drought ending, our plant will be tentatively designed to produce 15 million gallons of drinkable water per day to homes and businesses in the Los Angeles county area. With a recovery rate of near 30%, and a requirement for diluting any brine, the input feed flowrate will be 70.5 million gallons per day. Following the similarly built Carlsbad plant, we expect the plant to last 30 years. The project building timeline is one year based on our relatively small plant size. We will run for 350 days a year while shutting down for cleaning and maintenance once every 3 months (based on DOW standards).<ref name = "Bates">Bates, Wayne. “Cleaning your RO”. Hydranautics. </ref> The plant will be built 500 meters from the coastline on Long Beach. A borehole will be drilled 10 meters into the ground. At this depth there should be seawater aquifers below the coast that regularly recharges itself with seawater from the ocean. A vertical well containing intake holes will be put into the borehole. The intake holes will be equipped with large size filters that can easily be replaced. At the surface, the well will be connected to a large-scale pump in a well pump house. From there, the seawater will be horizontally pumped into a seawater tank in the desalination plant. This will be considered the start of our process. The process ends after pretreatment, desalination, post treatment, and brine treatment are completed. Appendix 1 shows more details on the Design Basis.

The feed composition is assumed to be the standard seawater “average” based on the DOW Chemical ROSA 9.0 Technical Manual in Table 1. However, in reality the seawater composition at the plant site will have to be chemically tested. The final product will be based on L.A. standards for drinking water quality (Table 1). The specifications in the table are upper limits that we cannot exceed. We will aim to reach the salt level standards but any remineralization and pH treatment will be exported and paid for due to the added complexity and cost in making such a system in-house.

Due to the acidity of the desalinated water coming out of the reverse osmosis (RO) system (pH ~ 6), the pumps and pipes carrying the desalinated water for post treatment will likely corrode over time. We have decided to place the desalination plant by the Long Beach Water Reclamation Plant (about 50 meters away) to prevent pipe corrosion.

[[File:Table_1.jpg|center|200px|frame|Table 1. Concentration of the standard seawater, the output from the RO system, and LA standard.]]

=Desalination Overview and Alternatives=

We first began our research into the current landscape of the desalination industry, particularly of plants that were built in the United States. We discovered there were many processes but the main ones included: 
* Membrane Separation (Forward Osmosis and Reverse Osmosis)
* Thermal Separation (Multi Stage Flash Distillation)
* Phase Change Separation (Freeze Separation)
Each process was analyzed for their pros and cons. For our project design, we decided to go with reverse osmosis as our desalination process. This is because it is the most commonly used process in the industry, and for good reason. Unlike thermal separation, which requires huge amounts of energy to supply heat, reverse osmosis is very energy efficient. This reduces energy costs that would be spread across the plant’s lifetime.<ref name = "Republic">Republic of Mauritius NY. "Technical Aspects of Desalination Plant". Sustainable Sanitation and Water Management. </ref>

Forward osmosis is a newer technology being developed that also uses membranes to separate water and salt and may be more energy efficient overall than reverse osmosis. This is because forward osmosis doesn’t require a pump to pressurize seawater through the membranes, but rather leaves a higher salt concentrated solution (AKA draw solution) on the other side of the membrane for the seawater to be drawn into. Then a second step is required to regenerate the concentrated draw solution and produce purified water.<ref>Mallinson, Alissa. "Study Shows Forward Osmosis Desalination Not Energy Efficient." MIT News [Boston]. July 23, 2014.</ref> We decided to move away from forward osmosis because it is a process that is still being refined and developed for desalination and lacks the data, programs, and technologies that reverse osmosis users benefit from. Additionally, a recent study from MIT has claimed that due to the inherent nature of forward osmosis requiring that seawater be drawn to a higher concentrated salt solution before the water and salt is separated, it cannot possibly be more energy efficient than reverse osmosis.<ref>O'Hern, Sean C., Karnik, Rohit. Nanofiltration across Defect-Sealed Nanoporous Monolayer Graphene. American Chemical Society 15.5 (2014); 3254-260. </ref>

Other processes such as freeze separation were too restrictive in their feed and process demands and required a lot of capital. Ultimately, reverse osmosis is simply the best process that exists for desalination and has been ranked the best when performance is evaluated in terms of technical, environmental, and economic aspects. It is a process that is flexible in water quantity, operation start-up and shut-off, and site location.

Nevertheless there are disadvantages to reverse osmosis that must be considered. The feed seawater must not be too salty and too low quality. The pressure to pump the seawater is enormous. The construction time is long for large plants. The capital costs are also high when compared with other processes such as thermal separation. But all of these disadvantages are ultimately outweighed by the advantages reverse osmosis provides.<ref name = "Republic"/>

=Process Overview and Alternatives=

See Appendix 2 (Figure 1) for the Process Flow Diagram (PFD). The process itself is detailed below. The stream flow rates are in Appendix 2 (Table 1). We assumed 10 meters of pipe for every stream except the inlet and outlet streams of the entire PFD.

==Pretreatment==

After the 10 meter vertical well pumps seawater from the Long Beach coast, the feed is pumped 500 meters to a seawater tank that makes it easy to shut-down the system and reduce the large volume of flow. The level control keeps the seawater from overfilling the tank. From there, part of the seawater goes to brine treatment while the rest go to pretreatment. The purpose of pretreatment in the desalination process is to prevent equipment damage by removing particles and microorganisms from the water. A variety of pollutants -- including calcium carbonate, sulfates, silica, clay, bacteria, and other colloids and microorganisms -- can disrupt the reverse osmosis process. Media filtration uses fine grains such as sand or anthracite to filter the particles. We used a large filter to catch larger particle material (e.g. sand). However, simple filtration is not enough to prevent colloid fouling. An ultrafilter (i.e. microfiltration) is used to remove virtually all suspended matter, effectively preventing colloid fouling. The costs associated with ultrafiltration are not prohibitive in terms of the project’s budget.<ref name = "Desalination">Desalination Pre-treatment. Lenntech.</ref>

Reverse osmosis systems can also be disrupted by the growth of microbiological films of bacteria on the film and pipe walls. They are typically treated as colloids during pretreatment, and most organisms are removed from the water by ultrafiltration, but this does not remove the smallest of organic particles. DBNPA is the most effective biocide, but these aren’t often used and have to be removed from the water. Chlorination is an effective way to remove bacteria, but it can cause degradation of the membrane, and must be removed before the reverse osmosis occurs. In our process we will use UV irradiation to remove any organic material that passes through the microfilters.<ref name = "Desalination"/>

Scaling can reduce the flux across the membrane and can build up in pipe corners, causing backup. It is caused by inorganic salts -- specifically sulfates and calcium and magnesium carbonate -- present in the water. Scaling can be prevented and/or controlled with acids, antiscalants, or ion exchange softening. Antiscalants are good for removing sulfates, but not much else. Acidification will only be effective at removing CaCO3 from water but, the decreased pH may deteriorate the membrane and pipes, so it may require future research.<ref name = "Desalination"/> Ion exchange is very expensive and requires more capital to undergo, but is overall the more effective option.

The ion exchanger is a pretreatment system which acts as a method to fight membrane scaling in reverse osmosis. The ion exchanger works by using sodium carboxylate as a resin which will react with Chlorine and Magnesium ions, removing them from the feed. The sodium is substituted out and enters the feed stream, increasing the overall sodium level and TDS. The lack of magnesium and chlorine ions means that there is nothing to react with the carbonates to form scaling salts on the membranes.

==Desalination (Reverse Osmosis)==

Two pumps (P-105 & P-106) take the adjusted feed seawater up to 58 bar. However, before this pumping occurs, about 70% of the original stream gets separated and heads into a PX Pressure Exchanger. This separated stream later joins together with the pumped stream to enter the reverse osmosis system at the required pressure of 58 bar. The streams can combine together as they have equal pressure. It is the concentrate stream from the output of the reverse osmosis system that transfers its energy to the adjusted feed in the PX Pressure Exchanger.

[[File:Figure_3.jpg|left|300px|frame|Figure 3. Inside look of pressure exchanger]]
The pressure of the concentrate stream is about 57 bar and it is fed into an exchanger which works by using vessel that resembles a small pipe within a rotating wheel. The low pressure (LP) feed enters the vessel at 1 bar (see Figure 3). The feed turns and is pushed out into an outlet pipe by the high pressure (HP) concentrate stream at ~57 bar. The feed is now a HP stream via this exchanged energy. The now LP concentrate stream turns and is pushed out into an outlet pipe by the LP feed entering the vessel.

98% of the energy within the concentrate stream can be recovered via the PX Pressure Exchanger.<ref>Desalination Product and Applications Catalog 2013. “Advance Technologies for Seawater Reverse Osmosis and Brackish Water Applications”. Energy Recovery, Inc. 2012-2014.</ref >Thus, the now HP feed stream is not quite 58 bar and will need to be pumped by a small booster pump (P-107).

When designing a reverse osmosis system, we need to make decisions on what specific kind of reverse osmosis system will be implemented. When making the decision, we investigated the advantages of different types of reverse osmosis among the following options: batch vs. continuous process, number of stages, and number of elements.<ref>FILMTEC Reverse Osmosis Membranes. “Water & Process Solutions Technical Manual.” Dow Chemical Company.</ref> Our decision was made based on process optimization which is described later in this report. The end result of that optimization was to go with a single stage, continuous process with five elements in series.

=Post Treatment=

The permeate is sent to post treatment after the salt content is analyzed and controlled if too high or too low due to any unexpected errors. As stated above, the post treatment process will be done 50 meters from our plant by the Long Beach Water Reclamation Plant. The reason we aren’t performing post treatment “in-house” is due to the increased cost it would take for us to build and run the equipment for post treatment. We have the option to use the government’s water treatment facilities that already exist and our 15 million gallons per day output is a relatively small addition to the large Long Beach Water Reclamation Plant. We will assume any costs to post treatment will be pushed to the consumer in the price of purchasing the drinking water from the Long Beach Water Reclamation Plant.

Post treatment at any facility works to bring water to “taste standards”. The water that is the permeate of our process is already meeting L.A. drinking standards as shown above in Table 1 after our mass balance calculations around the reverse osmosis system (Appendix 3, Table 1). The post treatment increases the pH from 6.3 to 7.3 with the addition of base and then adds minerals like fluoride, calcium, and magnesium back in to improve taste.

==Brine Treatment and Byproduct Waste==

A natural consequence of a reverse osmosis procedure is that a lot of salt will be leftover from the membrane separation in a very concentrated form. This is called “brine” and will have to be taken care of. According to EPA requirements, the salt concentration of water entering the environment (via land or sea) cannot be higher than 40,000 ppm.8 The salt concentration of our brine is much higher (around 70,000 ppm). Thus, the brine must be removed or treated. We decided to treat the brine by diluting it as removing brine can be difficult and may require lots of land (e.g. evaporation ponds).

After the concentrate exchanges energy in the energy exchanger, it becomes a low pressure stream. It is pumped into a dilution tank and mixes with 18.2 million gallons per day of seawater from the inlet to become diluted brine which is pumped 600 meters out to sea.

The waste byproducts are the solid buildup on the filters and membranes. The filters and membranes also have a set lifetime and thus will also need to be removed. The waste is all non-hazardous and thus can be put in the municipal dump.

=Process Details and Equipment Sizing=

==Filters and UV Treatment [Pretreatment]==
The cost of our ultrafiltration system is estimated using data from the University of Tennessee at Knoxville.<ref>Qiang He and R. Bruce Robinson Membrane Filtration Systems. “Membrane Module Costs.” University of Tennessee at Knoxville. December 2014.</ref> It estimated the cost of the cost of the ultrafiltration $0.057 per 1000 gallons and the construction and capital costs are $0.39 per gallon per day. The cost of UV treatment is based off of units available from buyultraviolet.com for units that can treat about 450 gallons per minute.<ref>Atlantic Ultraviolet Corporation. “Megatron Water Disinfection System.”</ref> Due to the water having to be in relatively small pipes for treatment we do not expect the costs to scale very much, so we believe these units to be an accurate cost estimate.

==Ion Exchanger [Pretreatment] ==

The ion exchanger was sized using data from DOW. Our process will use 60 vessels with ion exchange beds of volume 5.73 m3. With a flow rate of 30 bed volumes/hour (10,312.2 m3/hr) the system of beds has a capacity of 125% of the feed. This allows for periodic regeneration of the resin, a weak acid cation resin (DOW MAC-3 Resin). This weak cation resin will remove cationic calcium and magnesium from the water to prevent scaling. The mass flow was calculated via a mass balance on the ion exchanger (Appendix 3).

==Reverse Osmosis System [Desalination]==

Each of the 2607 Reverse Osmosis SW30HRLE-440i [Seawater 30 High Rejection Low Energy 440 ft2 interlocking cap membrane] units have the same dimensions. A, the net length, is 1016 mm. B, the total length including the endcap, is 1029 mm. C, the diameter of the filter element, is 201mm. D, the diameter of the permeate pathway, is 29 mm. For a single stage, we have five of the elements in series so the overall length of a single stage is expected to be roughly 5145 mm. The membrane material is polyamide thin film composite. It will last three years.
Operating Conditions [ROSA 9.0]: 
* Temperature: 30oC 
* Pressure: 58 bar 
* Flow Rate: 52.3 Million gal per day 
* Recovery: 28.7%

[[File:Figure_4.png|center|400px|frame|Figure 4. Diagram of Dow FILMTEC SW30HRLE-440i]]
Cleaning Schedule: 
* Conservative Estimate: Once every three months (three days available for cleaning)<ref name = "Bates"/> 
* DOW FILMTEC Cleaning Conditions 
** The normalized permeate flow drops 10%
* The normalized salt passage increases 5 - 10%

==Seawater Holding Tank and Dilution Tank [Brine Treatment]==

The initial seawater holding tank works as a buffer to prevent surges and help perform shutdowns quickly and easily. We expect to run continuously for 350 days a year with the rest of the available time spread out and used for shutting down the plant for cleaning membranes, filters, and the fixing of unexpected errors. There is a big tradeoff between having a holding tank with high retention time and having a holding tank with low cost. With that in mind, we erred on the side of lowering capital cost and went with a retention time of one hour. The volume could be found by multiplying the retention time with the feed flow rate (Stream 1). We assumed that a good height to width ratio for a large water tank would be 4:1.

A mass balance was done around the dilution tank to find what amount of seawater would be necessary to dilute the brine to 45,000 ppm salinity (Appendix 3; Table 3). With this the feed flow rate into the tank could be found. It was found that it takes about 20 minutes of retention time for good dilution of the brine in seawater.8 We again assumed that a good height to width ratio for a large water tank would be 4:1.

We assumed a corrosion allowance of 4 mm wall thickness. We calculated wall thickness using vessel pressure, hoop stress, longitudinal stress, and weld factor. Each tank is made with SS316 to protect against corrosion.<ref>Towler G, Sinnott R. Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013.</ref>

<center>
'''Table 2.''' Tank sizes.
[[File:Table_2.jpg|center|600px]]
</center>

=Process Optimization=
<center> '''Table 3.'''Optimization Criteria
[[File:Table_3.jpg|center|400px]]
</center>
In order to optimize the reverse osmosis system, numerous simulations were run. Three representative cases are reported as shown above (Table 3). The objective of the optimization was to minimize the required water price to break even in 10 years. In order to achieve this goal, the balance between the capital cost from the number of the filter elements and utility cost from energy consumption was weighed. Before the energy recovery system was implemented, the general trend was that if we have more filter elements in series, overall recovery increases but recovery rate per filter decreases, and we need more filter elements. While increasing the elements in series increase the required number of filters, and thereby increases the capital cost, we require less energy to produce the same amount of water. After the energy recovery system is implemented, the correlation between the number of elements in series and the required energy becomes more complicated because high pressure concentrate is used to pressurize the feed. However, the trade-off between the utility cost and capital cost still exists. Of the three cases shown above, single stage, five elements RO system turned out to be the most optimal solution because single stage, single element system required higher utility cost and two stage system slightly did not meet the LA standard. Detailed information which was used for the optimization is available in Appendix 3 and in Appendix 6.

=Ethical, Environmental, and Societal Considerations=

There are a few key considerations we have followed to meet ethical, environmental, and societal guidelines and requirements: 
* Meet L.A. Drinking Water Standards
** Correct ion concentration
* Moved near water treatment plant
** Pipes will not corrode from low pH
* Working below pressure limits
* Water will be locally served and meet taste standards
** Via Post Treatment mineralization
* Follows EPA Environmental Salinity Requirements
* Reduced Energy Requirements
** PX Pressure Exchanger
** Optimized RO System
** Reverse Osmosis is itself Energy Saving

For further environmental considerations, see effluent summary in Appendix 4.

=Economic Analysis=

The pricing of the majority of the process was completed using the Aspen Economic Analyzer, however we had determine costs for the ultrafiltration, reverse osmosis membranes, and the ultraviolet treatment was completed separately. This led to a total ISBL of $56 MM, OSBL of $22.4 MM, Engineering Costs of $22.4 MM and a Contingency of $7.8 MM. We also chose a working Capital of $10 MM. This working capital is sufficient because we have no raw materials costs. We also had FCOP of $5.2 MM and VCOP of $3.5 MM.

We originally were planning on pricing our water at $700 per acre-foot, but after doing economic analysis we had a 10 year NPV of $-76.2 MM. After determining did a sensitivity analysis and determined that if our water was priced at $2104 per acre-foot we could have an NPV of $0 after 10 years and if we were able to get the same price of $2257 per acre-foot as the plant in Carlsbad we would have a 10 year NPV of $8.3 MM which increases to $41.1 MM at 20 years. For more detailed data about the economic analysis see Appendix 6.

=Conclusion and Recommendations=

Our mission was to create an economically viable process that will desalinate water for human consumption in the L.A. County. We have found after research and optimization that the best process to undergo desalination is a continuous single stage, five elements reverse osmosis system. Our process can meet the required L.A. drinking water standard and we have set up plans for post treatment at a local water reclamation site. While meeting the drinking water standards, we also met our ethical, environmental, and social standards.

Nevertheless, at a sale price of $700 per acre-foot, our proposed desalination plant in Los Angeles, California looks to be a very expensive project to undergo. Without any sort of government assistance to carry the heavy capital costs of such a project, we would suggest the plant not be built. However, after performing a sensitivity analysis at higher (and less conservative) price points we reached a breakeven point in 10 years with a sale price of $2104 per acre-foot of drinking water. Research has shown that the droughts striking California look only to get worse in the near future. With limited freshwater resources, there is no option other than to consider desalination. Therefore it is reasonable to expect we could get near the same price deal as local desalination plants like the Carlsbad Plant ($2257 per acre-foot) As we have met all design constraints for this project, we hope to convince policymakers to adopt our project at least $2104 per acre-foot water. If this deal is reached, we recommend that this project and design be funded.

However, there are ways to improve our design with future work. Right now all the pipes diameters are based on having a flowrate of 1.5 m/s. But in reality, we’d need to use standard piping sizes to get more accurate pricing and energy costs. Also, in reality we would be purchasing some of these vessels in bulk quantities (i.e. bulk filters) so the price could be lowered with negotiation and economies of scale. We also have pumps that are providing very little pressure for their high cost. We can remove these pumps and simply run at higher pressures earlier. We can potentially reduce early energy costs by building a “floating” desalination plant in the sea to reduce pumping distance from the well.

=Appendices=

==Appendix 1: Design Basis==
[[File:Appendix_1.png|thumb|border|center|800px|middle|]]

==Appendix 2: Process Flow Diagram==
[[File:Appendix_2.png|center]]
==Appendix 3: Mass and Energy Balance Tables==
'''Table 1'''. Mass balance on 1 stage 5 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T1.png|center]]

'''Table 2'''. Mass balance on 1 stage 1 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T2.png|center]]

'''Table 3'''. Mass balance on 2 stage 5,5 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T3.png|center]]

'''Table 4'''. 1 stage 5 elements CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T4.png|center]]

'''Table 5'''. 1 stage 1 element CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T5.png|center]]

'''Table 6'''. 2 stages 5,5 elements CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T6.png|center]]

'''Table 7'''. Energy balance on 1 stage 5 elements RO system
[[File:Appendix_3_T7.png|center]]

==Appendix 4: Effluent Summary==
[[File:Appendix_4.png|center]]

==Appendix 5: Cost of Production==
[[File:Appendix_5.png|center]]

==Appendix 6: Economic Anlaysis==
===Case I ===
Water is $700; RO with 1 stage and 5 elements
[[File:Appendix_6_C1.png|center]]

===Case II===
Water is $1400; RO with 1 stage and 5 elements
[[File:Appendix_6_C2.png|center]]

===Case III: Carlsbad ===
Water is $2257; RO with 1 stage and 5 elements
[[File:Appendix_6_C3.png|center]]

===Case IV: Breakeven (NPV = 0 at 10 years)===
Water is $2104; RO with 1 stage and 5 elements
[[File:Appendix_6_C4.png|center]]

===Case V: Breakeven (NPV = 0 at 10 years)===
Water is $2112; RO with 2 stages and 5,5 elements
[[File:Appendix_6_C5.png|center]]

===Case VI: Breakeven (NPV = 0 at 10 years)===
Water is $2138 with 1 stage and 1 element
[[File:Appendix_6_C6.png|center]]

==Appendix 7: ROSA Simulation Outputs==

'''Table 1'''. System detail of 1 stage 5 elements RO system
[[File:Appendix_7_T1.png|center]]

'''Table 2'''. System detail of 1 stage 1 elements RO system
[[File:Appendix_7_T2.png|center]]

'''Table 3'''. System detail of 2 stage 5,5 elements system
[[File:Appendix_7_T3.png|center]]

'''Table 4'''. 1 stage 5 elements RO system concentrations and pH in each stream
[[File:Appendix_7_T4.png|center]]

'''Table 5'''. 1 stage 1 element RO system concentration and pH in each stream
[[File:Appendix_7_T5.png|center]]

'''Table 6'''. 2 stage 5,5 elements RO system concentrations and pH in each stream
[[File:Appendix_7_T6.png|center]]

=References=
<references/>

Desalination - Team E

2016-03-12T03:25:13Z

Bts511: /* Appendix 1: Design Basis */

Group E Corporation

Authors: Hassan Ali, Woo Soo Choe, Brett Sleyster, Jake Stolley

Instructors: Fengqi You, David Wegerer

March 11, 2016

=Executive Summary=

Outlined in this report is the project model report for the proposed desalination plant in Los Angeles, California which will provide drinking water to local citizens using reverse osmosis process. This location was chosen due to the region’s propensity for droughts in recent years. The fresh water resources in California have plummeted as water prices have risen 3-4% each year. This is only expected to continue. Other companies have worked deals with the Californian government to create desalination plants that provide a new fresh water drinking source. This suggests that there is a strong market here with potential growth.

The plant will have an output of 15 million gallons per day, which is about medium size relative to other plants. Thus the build time will be about one year. The plant will run for 30 years, with 350 days per year of use. Shutdown time will be used for cleaning and maintenance of equipment. The plant will be built on Long Beach, 500 meters from the coast nearby the Long Beach Water Reclamation Plant (LBWP). The input will withdraw 70.5 million gallons per day of seawater that will be used for reverse osmosis and dilution of brine.

Reverse osmosis was chosen as the desalination method for its low energy cost and flexibility. We designed a process that will pretreat the seawater to remove waste, colloids, bacteria, and membrane-damaging scalants. Once this seawater is treated is pumped to heavy pressures through a reverse osmosis system. After doing process optimization under various system parameter changes, we found a continuous, single stage, five element reverse osmosis process to be best. It produces a near 30% recovery. We reduce energy costs by including a pressure exchanger between our concentrate outlet and adjusted feed into the reverse osmosis system.

After performing mass and energy balances, we were able to size equipment and find our permeate composition. It turns out our process will meet L.A. standards for drinking water quality. However, post treatment will be done by the local LBWP to make the water taste better and be non-caustic to pipes. This will reduce “in-house” costs. Many environmental considerations were made, to meet our environmental standards. We meet EPA salinity requirements for our brine waste by diluting it. We also don’t use a lot of energy and thus reduce our fossil fuel consumption and CO2 generation. To be safe, we work well under max pressure limits for our equipment. We make sure we are ethical and meet all standards that our local county requires of us.

After an economic evaluation at a conservative well-water price of $700 per acre-foot, we realize that our revenue is way too low to cover our costs. Capital costs are especially large with regards to our large water tanks. We find that we reach a 10 year NPV breakpoint of $2104 per acre-foot after performing a sensitivity analysis. The Carlsbad Plant in San Diego worked out a deal for $2257 per acre-foot.

As we have met all design constraints for this project, we hope to convince policymakers to adopt our project at least $2104 per acre-foot water. If this deal is reached, we recommend that this project and design be funded. There is future work to be done to improve costs such as performing pump optimization. However, this work is minimal and won’t harm our bottom line if we reach our proposed price points.

__TOC__

=Introduction=

Due to a combination of various factors--including rising human population, climate change, increased energy consumption, and land erosion--the availability of fresh water as a resource for human consumption has been dwindling. This has had a negative impact on the world, but this impact has been especially felt in areas with natural shortages in fresh water. Desert regions, such as Saudi Arabia, have resorted to desalinating seawater at the coast and transporting up to their cities. This encompasses their entire freshwater system as very few natural water sources remain. In the United States there is a region which has shown increasing signs of following in the footsteps of Saudi Arabia and becoming fully reliant on desalination: Los Angeles, California.

=Location and Market Analysis=

In this project, Los Angeles was chosen as the location our plant, because the area has been subject to a severe drought for the last four years, as shown in Figure 1. The rapid decrease in the available water has caused permanent damages to the ground, decreased the amount of water in the reservoir from 45% to 25%, and reduced the amount of hydroelectric power produced in the state. Such deficiency in usable water has increased the water prices in Los Angeles by 6-7% annually for the last decade and is expected to rise by 3.4% annually for typical customers for the next five years. For customers who use more than 20000 gallons a year, the bill is expected to increase by 34% by 2021. This price increase is occurring because of both drought and maintenance. <ref>Pacific Institute: The California Drought. 2014.</ref>

[[File:Figure_1.jpg|right|200px|frame|Figure 1. California Drought Index.]]
As of March 2015 the Metropolitan Water District of Southern California was willing to pay as much as $700 per acre-foot of water to farmers in the Sacramento Valley, so we assume this will be the target price of the water produced by the this desalination plant.<ref>Hessel, Phil. “Dry Southern California Offers Northern Farmers Top Dollar for Water”. NBC News. March 18, 2015. </ref>

Currently, there is a plant in Carlsbad which supplies drinking water to the San Diego area. The plant can produce up to 50 million gallons per day and is the nation's largest water desalination plant. The water from this plant costs $2257 per acre-foot for the first 48000 acre-feet and about $2000 for any additional acre-foot. The county signed a contract with this plant to purchase the first 48000 acre-feet of water for the next 30 years.<ref>Carlsbad Desalination Plant. “Project Overview”. January 2016.</ref> Also, there is another desalination plant in Tampa that produces about 25 million gallons of water per day, and over 2000 other public and industrial plants with production capacity of over 300000 gallons per day. When conducting the economic analysis our plant, such numbers presented by other plants must be considered, so our project can yield positive profit.<ref>Leven, Rachel. “U.S. Desalination Industry Grows Since 2000; See as Essential to Meeting Supply Needs”. Bloomberg BNA. August 21, 2013.</ref>

The profitability of the project is naturally heavily dependent on the price of water in California. If the drought ends, water supply will increase and the water price will drop significantly, but if the drought continues, the price of water in California could rise even more rapidly and allow desalination plants to be more profitable.<ref>Southern California Public Radio. “Where is California water use decreasing?”. December 2015.</ref> It may be difficult to make this plant profitable without a guaranteed purchase or other subsidy from the local governments because the presence of pre-existing plants maintains the water price relatively low. However, external sources of profit such as guaranteed purchase or subsidy are expected because the ramifications of not having desalination plant or other sources of water may be significantly more expensive than the subsidizing a desalination plant assuming the drought continues.<ref>Weiser, Matt. “Could desalination solve California’s water problem?”. The Sacramento Bee. October 18, 2014.</ref>

=Design Basis and Technical Approach=

[[File:Figure_2.jpg|left|200px|frame|Figure 2. Vertical Well Feed]]
Considering the number of pre-existing desalination plant and the potential risk of the drought ending, our plant will be tentatively designed to produce 15 million gallons of drinkable water per day to homes and businesses in the Los Angeles county area. With a recovery rate of near 30%, and a requirement for diluting any brine, the input feed flowrate will be 70.5 million gallons per day. Following the similarly built Carlsbad plant, we expect the plant to last 30 years. The project building timeline is one year based on our relatively small plant size. We will run for 350 days a year while shutting down for cleaning and maintenance once every 3 months (based on DOW standards).<ref name = "Bates">Bates, Wayne. “Cleaning your RO”. Hydranautics. </ref> The plant will be built 500 meters from the coastline on Long Beach. A borehole will be drilled 10 meters into the ground. At this depth there should be seawater aquifers below the coast that regularly recharges itself with seawater from the ocean. A vertical well containing intake holes will be put into the borehole. The intake holes will be equipped with large size filters that can easily be replaced. At the surface, the well will be connected to a large-scale pump in a well pump house. From there, the seawater will be horizontally pumped into a seawater tank in the desalination plant. This will be considered the start of our process. The process ends after pretreatment, desalination, post treatment, and brine treatment are completed. Appendix 1 shows more details on the Design Basis.

The feed composition is assumed to be the standard seawater “average” based on the DOW Chemical ROSA 9.0 Technical Manual in Table 1. However, in reality the seawater composition at the plant site will have to be chemically tested. The final product will be based on L.A. standards for drinking water quality (Table 1). The specifications in the table are upper limits that we cannot exceed. We will aim to reach the salt level standards but any remineralization and pH treatment will be exported and paid for due to the added complexity and cost in making such a system in-house.

Due to the acidity of the desalinated water coming out of the reverse osmosis (RO) system (pH ~ 6), the pumps and pipes carrying the desalinated water for post treatment will likely corrode over time. We have decided to place the desalination plant by the Long Beach Water Reclamation Plant (about 50 meters away) to prevent pipe corrosion.

[[File:Table_1.jpg|center|200px|frame|Table 1. Concentration of the standard seawater, the output from the RO system, and LA standard.]]

=Desalination Overview and Alternatives=

We first began our research into the current landscape of the desalination industry, particularly of plants that were built in the United States. We discovered there were many processes but the main ones included: 
* Membrane Separation (Forward Osmosis and Reverse Osmosis)
* Thermal Separation (Multi Stage Flash Distillation)
* Phase Change Separation (Freeze Separation)
Each process was analyzed for their pros and cons. For our project design, we decided to go with reverse osmosis as our desalination process. This is because it is the most commonly used process in the industry, and for good reason. Unlike thermal separation, which requires huge amounts of energy to supply heat, reverse osmosis is very energy efficient. This reduces energy costs that would be spread across the plant’s lifetime.<ref name = "Republic">Republic of Mauritius NY. "Technical Aspects of Desalination Plant". Sustainable Sanitation and Water Management. </ref>

Forward osmosis is a newer technology being developed that also uses membranes to separate water and salt and may be more energy efficient overall than reverse osmosis. This is because forward osmosis doesn’t require a pump to pressurize seawater through the membranes, but rather leaves a higher salt concentrated solution (AKA draw solution) on the other side of the membrane for the seawater to be drawn into. Then a second step is required to regenerate the concentrated draw solution and produce purified water.<ref>Mallinson, Alissa. "Study Shows Forward Osmosis Desalination Not Energy Efficient." MIT News [Boston]. July 23, 2014.</ref> We decided to move away from forward osmosis because it is a process that is still being refined and developed for desalination and lacks the data, programs, and technologies that reverse osmosis users benefit from. Additionally, a recent study from MIT has claimed that due to the inherent nature of forward osmosis requiring that seawater be drawn to a higher concentrated salt solution before the water and salt is separated, it cannot possibly be more energy efficient than reverse osmosis.<ref>O'Hern, Sean C., Karnik, Rohit. Nanofiltration across Defect-Sealed Nanoporous Monolayer Graphene. American Chemical Society 15.5 (2014); 3254-260. </ref>

Other processes such as freeze separation were too restrictive in their feed and process demands and required a lot of capital. Ultimately, reverse osmosis is simply the best process that exists for desalination and has been ranked the best when performance is evaluated in terms of technical, environmental, and economic aspects. It is a process that is flexible in water quantity, operation start-up and shut-off, and site location.

Nevertheless there are disadvantages to reverse osmosis that must be considered. The feed seawater must not be too salty and too low quality. The pressure to pump the seawater is enormous. The construction time is long for large plants. The capital costs are also high when compared with other processes such as thermal separation. But all of these disadvantages are ultimately outweighed by the advantages reverse osmosis provides.<ref name = "Republic"/>

=Process Overview and Alternatives=

See Appendix 2 (Figure 1) for the Process Flow Diagram (PFD). The process itself is detailed below. The stream flow rates are in Appendix 2 (Table 1). We assumed 10 meters of pipe for every stream except the inlet and outlet streams of the entire PFD.

==Pretreatment==

After the 10 meter vertical well pumps seawater from the Long Beach coast, the feed is pumped 500 meters to a seawater tank that makes it easy to shut-down the system and reduce the large volume of flow. The level control keeps the seawater from overfilling the tank. From there, part of the seawater goes to brine treatment while the rest go to pretreatment. The purpose of pretreatment in the desalination process is to prevent equipment damage by removing particles and microorganisms from the water. A variety of pollutants -- including calcium carbonate, sulfates, silica, clay, bacteria, and other colloids and microorganisms -- can disrupt the reverse osmosis process. Media filtration uses fine grains such as sand or anthracite to filter the particles. We used a large filter to catch larger particle material (e.g. sand). However, simple filtration is not enough to prevent colloid fouling. An ultrafilter (i.e. microfiltration) is used to remove virtually all suspended matter, effectively preventing colloid fouling. The costs associated with ultrafiltration are not prohibitive in terms of the project’s budget.<ref name = "Desalination">Desalination Pre-treatment. Lenntech.</ref>

Reverse osmosis systems can also be disrupted by the growth of microbiological films of bacteria on the film and pipe walls. They are typically treated as colloids during pretreatment, and most organisms are removed from the water by ultrafiltration, but this does not remove the smallest of organic particles. DBNPA is the most effective biocide, but these aren’t often used and have to be removed from the water. Chlorination is an effective way to remove bacteria, but it can cause degradation of the membrane, and must be removed before the reverse osmosis occurs. In our process we will use UV irradiation to remove any organic material that passes through the microfilters.<ref name = "Desalination"/>

Scaling can reduce the flux across the membrane and can build up in pipe corners, causing backup. It is caused by inorganic salts -- specifically sulfates and calcium and magnesium carbonate -- present in the water. Scaling can be prevented and/or controlled with acids, antiscalants, or ion exchange softening. Antiscalants are good for removing sulfates, but not much else. Acidification will only be effective at removing CaCO3 from water but, the decreased pH may deteriorate the membrane and pipes, so it may require future research.<ref name = "Desalination"/> Ion exchange is very expensive and requires more capital to undergo, but is overall the more effective option.

The ion exchanger is a pretreatment system which acts as a method to fight membrane scaling in reverse osmosis. The ion exchanger works by using sodium carboxylate as a resin which will react with Chlorine and Magnesium ions, removing them from the feed. The sodium is substituted out and enters the feed stream, increasing the overall sodium level and TDS. The lack of magnesium and chlorine ions means that there is nothing to react with the carbonates to form scaling salts on the membranes.

==Desalination (Reverse Osmosis)==

Two pumps (P-105 & P-106) take the adjusted feed seawater up to 58 bar. However, before this pumping occurs, about 70% of the original stream gets separated and heads into a PX Pressure Exchanger. This separated stream later joins together with the pumped stream to enter the reverse osmosis system at the required pressure of 58 bar. The streams can combine together as they have equal pressure. It is the concentrate stream from the output of the reverse osmosis system that transfers its energy to the adjusted feed in the PX Pressure Exchanger.

[[File:Figure_3.jpg|left|300px|frame|Figure 3. Inside look of pressure exchanger]]
The pressure of the concentrate stream is about 57 bar and it is fed into an exchanger which works by using vessel that resembles a small pipe within a rotating wheel. The low pressure (LP) feed enters the vessel at 1 bar (see Figure 3). The feed turns and is pushed out into an outlet pipe by the high pressure (HP) concentrate stream at ~57 bar. The feed is now a HP stream via this exchanged energy. The now LP concentrate stream turns and is pushed out into an outlet pipe by the LP feed entering the vessel.

98% of the energy within the concentrate stream can be recovered via the PX Pressure Exchanger.<ref>Desalination Product and Applications Catalog 2013. “Advance Technologies for Seawater Reverse Osmosis and Brackish Water Applications”. Energy Recovery, Inc. 2012-2014.</ref >Thus, the now HP feed stream is not quite 58 bar and will need to be pumped by a small booster pump (P-107).

When designing a reverse osmosis system, we need to make decisions on what specific kind of reverse osmosis system will be implemented. When making the decision, we investigated the advantages of different types of reverse osmosis among the following options: batch vs. continuous process, number of stages, and number of elements.<ref>FILMTEC Reverse Osmosis Membranes. “Water & Process Solutions Technical Manual.” Dow Chemical Company.</ref> Our decision was made based on process optimization which is described later in this report. The end result of that optimization was to go with a single stage, continuous process with five elements in series.

=Post Treatment=

The permeate is sent to post treatment after the salt content is analyzed and controlled if too high or too low due to any unexpected errors. As stated above, the post treatment process will be done 50 meters from our plant by the Long Beach Water Reclamation Plant. The reason we aren’t performing post treatment “in-house” is due to the increased cost it would take for us to build and run the equipment for post treatment. We have the option to use the government’s water treatment facilities that already exist and our 15 million gallons per day output is a relatively small addition to the large Long Beach Water Reclamation Plant. We will assume any costs to post treatment will be pushed to the consumer in the price of purchasing the drinking water from the Long Beach Water Reclamation Plant.

Post treatment at any facility works to bring water to “taste standards”. The water that is the permeate of our process is already meeting L.A. drinking standards as shown above in Table 1 after our mass balance calculations around the reverse osmosis system (Appendix 3, Table 1). The post treatment increases the pH from 6.3 to 7.3 with the addition of base and then adds minerals like fluoride, calcium, and magnesium back in to improve taste.

==Brine Treatment and Byproduct Waste==

A natural consequence of a reverse osmosis procedure is that a lot of salt will be leftover from the membrane separation in a very concentrated form. This is called “brine” and will have to be taken care of. According to EPA requirements, the salt concentration of water entering the environment (via land or sea) cannot be higher than 40,000 ppm.8 The salt concentration of our brine is much higher (around 70,000 ppm). Thus, the brine must be removed or treated. We decided to treat the brine by diluting it as removing brine can be difficult and may require lots of land (e.g. evaporation ponds).

After the concentrate exchanges energy in the energy exchanger, it becomes a low pressure stream. It is pumped into a dilution tank and mixes with 18.2 million gallons per day of seawater from the inlet to become diluted brine which is pumped 600 meters out to sea.

The waste byproducts are the solid buildup on the filters and membranes. The filters and membranes also have a set lifetime and thus will also need to be removed. The waste is all non-hazardous and thus can be put in the municipal dump.

=Process Details and Equipment Sizing=

==Filters and UV Treatment [Pretreatment]==
The cost of our ultrafiltration system is estimated using data from the University of Tennessee at Knoxville.<ref>Qiang He and R. Bruce Robinson Membrane Filtration Systems. “Membrane Module Costs.” University of Tennessee at Knoxville. December 2014.</ref> It estimated the cost of the cost of the ultrafiltration $0.057 per 1000 gallons and the construction and capital costs are $0.39 per gallon per day. The cost of UV treatment is based off of units available from buyultraviolet.com for units that can treat about 450 gallons per minute.<ref>Atlantic Ultraviolet Corporation. “Megatron Water Disinfection System.”</ref> Due to the water having to be in relatively small pipes for treatment we do not expect the costs to scale very much, so we believe these units to be an accurate cost estimate.

==Ion Exchanger [Pretreatment] ==

The ion exchanger was sized using data from DOW. Our process will use 60 vessels with ion exchange beds of volume 5.73 m3. With a flow rate of 30 bed volumes/hour (10,312.2 m3/hr) the system of beds has a capacity of 125% of the feed. This allows for periodic regeneration of the resin, a weak acid cation resin (DOW MAC-3 Resin). This weak cation resin will remove cationic calcium and magnesium from the water to prevent scaling. The mass flow was calculated via a mass balance on the ion exchanger (Appendix 3).

==Reverse Osmosis System [Desalination]==

Each of the 2607 Reverse Osmosis SW30HRLE-440i [Seawater 30 High Rejection Low Energy 440 ft2 interlocking cap membrane] units have the same dimensions. A, the net length, is 1016 mm. B, the total length including the endcap, is 1029 mm. C, the diameter of the filter element, is 201mm. D, the diameter of the permeate pathway, is 29 mm. For a single stage, we have five of the elements in series so the overall length of a single stage is expected to be roughly 5145 mm. The membrane material is polyamide thin film composite. It will last three years.
Operating Conditions [ROSA 9.0]: 
* Temperature: 30oC 
* Pressure: 58 bar 
* Flow Rate: 52.3 Million gal per day 
* Recovery: 28.7%

[[File:Figure_4.png|center|400px|frame|Figure 4. Diagram of Dow FILMTEC SW30HRLE-440i]]
Cleaning Schedule: 
* Conservative Estimate: Once every three months (three days available for cleaning)<ref name = "Bates"/> 
* DOW FILMTEC Cleaning Conditions 
** The normalized permeate flow drops 10%
* The normalized salt passage increases 5 - 10%

==Seawater Holding Tank and Dilution Tank [Brine Treatment]==

The initial seawater holding tank works as a buffer to prevent surges and help perform shutdowns quickly and easily. We expect to run continuously for 350 days a year with the rest of the available time spread out and used for shutting down the plant for cleaning membranes, filters, and the fixing of unexpected errors. There is a big tradeoff between having a holding tank with high retention time and having a holding tank with low cost. With that in mind, we erred on the side of lowering capital cost and went with a retention time of one hour. The volume could be found by multiplying the retention time with the feed flow rate (Stream 1). We assumed that a good height to width ratio for a large water tank would be 4:1.

A mass balance was done around the dilution tank to find what amount of seawater would be necessary to dilute the brine to 45,000 ppm salinity (Appendix 3; Table 3). With this the feed flow rate into the tank could be found. It was found that it takes about 20 minutes of retention time for good dilution of the brine in seawater.8 We again assumed that a good height to width ratio for a large water tank would be 4:1.

We assumed a corrosion allowance of 4 mm wall thickness. We calculated wall thickness using vessel pressure, hoop stress, longitudinal stress, and weld factor. Each tank is made with SS316 to protect against corrosion.<ref>Towler G, Sinnott R. Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013.</ref>

<center>
'''Table 2.''' Tank sizes.
[[File:Table_2.jpg|center|600px]]
</center>

=Process Optimization=
<center> '''Table 3.'''Optimization Criteria
[[File:Table_3.jpg|center|400px]]
</center>
In order to optimize the reverse osmosis system, numerous simulations were run. Three representative cases are reported as shown above (Table 3). The objective of the optimization was to minimize the required water price to break even in 10 years. In order to achieve this goal, the balance between the capital cost from the number of the filter elements and utility cost from energy consumption was weighed. Before the energy recovery system was implemented, the general trend was that if we have more filter elements in series, overall recovery increases but recovery rate per filter decreases, and we need more filter elements. While increasing the elements in series increase the required number of filters, and thereby increases the capital cost, we require less energy to produce the same amount of water. After the energy recovery system is implemented, the correlation between the number of elements in series and the required energy becomes more complicated because high pressure concentrate is used to pressurize the feed. However, the trade-off between the utility cost and capital cost still exists. Of the three cases shown above, single stage, five elements RO system turned out to be the most optimal solution because single stage, single element system required higher utility cost and two stage system slightly did not meet the LA standard. Detailed information which was used for the optimization is available in Appendix 3 and in Appendix 6.

=Ethical, Environmental, and Societal Considerations=

There are a few key considerations we have followed to meet ethical, environmental, and societal guidelines and requirements: 
* Meet L.A. Drinking Water Standards
** Correct ion concentration
* Moved near water treatment plant
** Pipes will not corrode from low pH
* Working below pressure limits
* Water will be locally served and meet taste standards
** Via Post Treatment mineralization
* Follows EPA Environmental Salinity Requirements
* Reduced Energy Requirements
** PX Pressure Exchanger
** Optimized RO System
** Reverse Osmosis is itself Energy Saving

For further environmental considerations, see effluent summary in Appendix 4.

=Economic Analysis=

The pricing of the majority of the process was completed using the Aspen Economic Analyzer, however we had determine costs for the ultrafiltration, reverse osmosis membranes, and the ultraviolet treatment was completed separately. This led to a total ISBL of $56 MM, OSBL of $22.4 MM, Engineering Costs of $22.4 MM and a Contingency of $7.8 MM. We also chose a working Capital of $10 MM. This working capital is sufficient because we have no raw materials costs. We also had FCOP of $5.2 MM and VCOP of $3.5 MM.

We originally were planning on pricing our water at $700 per acre-foot, but after doing economic analysis we had a 10 year NPV of $-76.2 MM. After determining did a sensitivity analysis and determined that if our water was priced at $2104 per acre-foot we could have an NPV of $0 after 10 years and if we were able to get the same price of $2257 per acre-foot as the plant in Carlsbad we would have a 10 year NPV of $8.3 MM which increases to $41.1 MM at 20 years. For more detailed data about the economic analysis see Appendix 6.

=Conclusion and Recommendations=

Our mission was to create an economically viable process that will desalinate water for human consumption in the L.A. County. We have found after research and optimization that the best process to undergo desalination is a continuous single stage, five elements reverse osmosis system. Our process can meet the required L.A. drinking water standard and we have set up plans for post treatment at a local water reclamation site. While meeting the drinking water standards, we also met our ethical, environmental, and social standards.

Nevertheless, at a sale price of $700 per acre-foot, our proposed desalination plant in Los Angeles, California looks to be a very expensive project to undergo. Without any sort of government assistance to carry the heavy capital costs of such a project, we would suggest the plant not be built. However, after performing a sensitivity analysis at higher (and less conservative) price points we reached a breakeven point in 10 years with a sale price of $2104 per acre-foot of drinking water. Research has shown that the droughts striking California look only to get worse in the near future. With limited freshwater resources, there is no option other than to consider desalination. Therefore it is reasonable to expect we could get near the same price deal as local desalination plants like the Carlsbad Plant ($2257 per acre-foot) As we have met all design constraints for this project, we hope to convince policymakers to adopt our project at least $2104 per acre-foot water. If this deal is reached, we recommend that this project and design be funded.

However, there are ways to improve our design with future work. Right now all the pipes diameters are based on having a flowrate of 1.5 m/s. But in reality, we’d need to use standard piping sizes to get more accurate pricing and energy costs. Also, in reality we would be purchasing some of these vessels in bulk quantities (i.e. bulk filters) so the price could be lowered with negotiation and economies of scale. We also have pumps that are providing very little pressure for their high cost. We can remove these pumps and simply run at higher pressures earlier. We can potentially reduce early energy costs by building a “floating” desalination plant in the sea to reduce pumping distance from the well.

=Appendices=

==Appendix 1: Design Basis==
[[File:Appendix_1.png|thumb|border|center|middle|800*800]]

==Appendix 2: Process Flow Diagram==
[[File:Appendix_2.png|center]]
==Appendix 3: Mass and Energy Balance Tables==
'''Table 1'''. Mass balance on 1 stage 5 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T1.png|center]]

'''Table 2'''. Mass balance on 1 stage 1 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T2.png|center]]

'''Table 3'''. Mass balance on 2 stage 5,5 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T3.png|center]]

'''Table 4'''. 1 stage 5 elements CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T4.png|center]]

'''Table 5'''. 1 stage 1 element CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T5.png|center]]

'''Table 6'''. 2 stages 5,5 elements CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T6.png|center]]

'''Table 7'''. Energy balance on 1 stage 5 elements RO system
[[File:Appendix_3_T7.png|center]]

==Appendix 4: Effluent Summary==
[[File:Appendix_4.png|center]]

==Appendix 5: Cost of Production==
[[File:Appendix_5.png|center]]

==Appendix 6: Economic Anlaysis==
===Case I ===
Water is $700; RO with 1 stage and 5 elements
[[File:Appendix_6_C1.png|center]]

===Case II===
Water is $1400; RO with 1 stage and 5 elements
[[File:Appendix_6_C2.png|center]]

===Case III: Carlsbad ===
Water is $2257; RO with 1 stage and 5 elements
[[File:Appendix_6_C3.png|center]]

===Case IV: Breakeven (NPV = 0 at 10 years)===
Water is $2104; RO with 1 stage and 5 elements
[[File:Appendix_6_C4.png|center]]

===Case V: Breakeven (NPV = 0 at 10 years)===
Water is $2112; RO with 2 stages and 5,5 elements
[[File:Appendix_6_C5.png|center]]

===Case VI: Breakeven (NPV = 0 at 10 years)===
Water is $2138 with 1 stage and 1 element
[[File:Appendix_6_C6.png|center]]

==Appendix 7: ROSA Simulation Outputs==

'''Table 1'''. System detail of 1 stage 5 elements RO system
[[File:Appendix_7_T1.png|center]]

'''Table 2'''. System detail of 1 stage 1 elements RO system
[[File:Appendix_7_T2.png|center]]

'''Table 3'''. System detail of 2 stage 5,5 elements system
[[File:Appendix_7_T3.png|center]]

'''Table 4'''. 1 stage 5 elements RO system concentrations and pH in each stream
[[File:Appendix_7_T4.png|center]]

'''Table 5'''. 1 stage 1 element RO system concentration and pH in each stream
[[File:Appendix_7_T5.png|center]]

'''Table 6'''. 2 stage 5,5 elements RO system concentrations and pH in each stream
[[File:Appendix_7_T6.png|center]]

=References=
<references/>

Desalination - Team E

2016-03-12T03:25:02Z

Bts511: /* Appendix 1: Design Basis */

Group E Corporation

Authors: Hassan Ali, Woo Soo Choe, Brett Sleyster, Jake Stolley

Instructors: Fengqi You, David Wegerer

March 11, 2016

=Executive Summary=

Outlined in this report is the project model report for the proposed desalination plant in Los Angeles, California which will provide drinking water to local citizens using reverse osmosis process. This location was chosen due to the region’s propensity for droughts in recent years. The fresh water resources in California have plummeted as water prices have risen 3-4% each year. This is only expected to continue. Other companies have worked deals with the Californian government to create desalination plants that provide a new fresh water drinking source. This suggests that there is a strong market here with potential growth.

The plant will have an output of 15 million gallons per day, which is about medium size relative to other plants. Thus the build time will be about one year. The plant will run for 30 years, with 350 days per year of use. Shutdown time will be used for cleaning and maintenance of equipment. The plant will be built on Long Beach, 500 meters from the coast nearby the Long Beach Water Reclamation Plant (LBWP). The input will withdraw 70.5 million gallons per day of seawater that will be used for reverse osmosis and dilution of brine.

Reverse osmosis was chosen as the desalination method for its low energy cost and flexibility. We designed a process that will pretreat the seawater to remove waste, colloids, bacteria, and membrane-damaging scalants. Once this seawater is treated is pumped to heavy pressures through a reverse osmosis system. After doing process optimization under various system parameter changes, we found a continuous, single stage, five element reverse osmosis process to be best. It produces a near 30% recovery. We reduce energy costs by including a pressure exchanger between our concentrate outlet and adjusted feed into the reverse osmosis system.

After performing mass and energy balances, we were able to size equipment and find our permeate composition. It turns out our process will meet L.A. standards for drinking water quality. However, post treatment will be done by the local LBWP to make the water taste better and be non-caustic to pipes. This will reduce “in-house” costs. Many environmental considerations were made, to meet our environmental standards. We meet EPA salinity requirements for our brine waste by diluting it. We also don’t use a lot of energy and thus reduce our fossil fuel consumption and CO2 generation. To be safe, we work well under max pressure limits for our equipment. We make sure we are ethical and meet all standards that our local county requires of us.

After an economic evaluation at a conservative well-water price of $700 per acre-foot, we realize that our revenue is way too low to cover our costs. Capital costs are especially large with regards to our large water tanks. We find that we reach a 10 year NPV breakpoint of $2104 per acre-foot after performing a sensitivity analysis. The Carlsbad Plant in San Diego worked out a deal for $2257 per acre-foot.

As we have met all design constraints for this project, we hope to convince policymakers to adopt our project at least $2104 per acre-foot water. If this deal is reached, we recommend that this project and design be funded. There is future work to be done to improve costs such as performing pump optimization. However, this work is minimal and won’t harm our bottom line if we reach our proposed price points.

__TOC__

=Introduction=

Due to a combination of various factors--including rising human population, climate change, increased energy consumption, and land erosion--the availability of fresh water as a resource for human consumption has been dwindling. This has had a negative impact on the world, but this impact has been especially felt in areas with natural shortages in fresh water. Desert regions, such as Saudi Arabia, have resorted to desalinating seawater at the coast and transporting up to their cities. This encompasses their entire freshwater system as very few natural water sources remain. In the United States there is a region which has shown increasing signs of following in the footsteps of Saudi Arabia and becoming fully reliant on desalination: Los Angeles, California.

=Location and Market Analysis=

In this project, Los Angeles was chosen as the location our plant, because the area has been subject to a severe drought for the last four years, as shown in Figure 1. The rapid decrease in the available water has caused permanent damages to the ground, decreased the amount of water in the reservoir from 45% to 25%, and reduced the amount of hydroelectric power produced in the state. Such deficiency in usable water has increased the water prices in Los Angeles by 6-7% annually for the last decade and is expected to rise by 3.4% annually for typical customers for the next five years. For customers who use more than 20000 gallons a year, the bill is expected to increase by 34% by 2021. This price increase is occurring because of both drought and maintenance. <ref>Pacific Institute: The California Drought. 2014.</ref>

[[File:Figure_1.jpg|right|200px|frame|Figure 1. California Drought Index.]]
As of March 2015 the Metropolitan Water District of Southern California was willing to pay as much as $700 per acre-foot of water to farmers in the Sacramento Valley, so we assume this will be the target price of the water produced by the this desalination plant.<ref>Hessel, Phil. “Dry Southern California Offers Northern Farmers Top Dollar for Water”. NBC News. March 18, 2015. </ref>

Currently, there is a plant in Carlsbad which supplies drinking water to the San Diego area. The plant can produce up to 50 million gallons per day and is the nation's largest water desalination plant. The water from this plant costs $2257 per acre-foot for the first 48000 acre-feet and about $2000 for any additional acre-foot. The county signed a contract with this plant to purchase the first 48000 acre-feet of water for the next 30 years.<ref>Carlsbad Desalination Plant. “Project Overview”. January 2016.</ref> Also, there is another desalination plant in Tampa that produces about 25 million gallons of water per day, and over 2000 other public and industrial plants with production capacity of over 300000 gallons per day. When conducting the economic analysis our plant, such numbers presented by other plants must be considered, so our project can yield positive profit.<ref>Leven, Rachel. “U.S. Desalination Industry Grows Since 2000; See as Essential to Meeting Supply Needs”. Bloomberg BNA. August 21, 2013.</ref>

The profitability of the project is naturally heavily dependent on the price of water in California. If the drought ends, water supply will increase and the water price will drop significantly, but if the drought continues, the price of water in California could rise even more rapidly and allow desalination plants to be more profitable.<ref>Southern California Public Radio. “Where is California water use decreasing?”. December 2015.</ref> It may be difficult to make this plant profitable without a guaranteed purchase or other subsidy from the local governments because the presence of pre-existing plants maintains the water price relatively low. However, external sources of profit such as guaranteed purchase or subsidy are expected because the ramifications of not having desalination plant or other sources of water may be significantly more expensive than the subsidizing a desalination plant assuming the drought continues.<ref>Weiser, Matt. “Could desalination solve California’s water problem?”. The Sacramento Bee. October 18, 2014.</ref>

=Design Basis and Technical Approach=

[[File:Figure_2.jpg|left|200px|frame|Figure 2. Vertical Well Feed]]
Considering the number of pre-existing desalination plant and the potential risk of the drought ending, our plant will be tentatively designed to produce 15 million gallons of drinkable water per day to homes and businesses in the Los Angeles county area. With a recovery rate of near 30%, and a requirement for diluting any brine, the input feed flowrate will be 70.5 million gallons per day. Following the similarly built Carlsbad plant, we expect the plant to last 30 years. The project building timeline is one year based on our relatively small plant size. We will run for 350 days a year while shutting down for cleaning and maintenance once every 3 months (based on DOW standards).<ref name = "Bates">Bates, Wayne. “Cleaning your RO”. Hydranautics. </ref> The plant will be built 500 meters from the coastline on Long Beach. A borehole will be drilled 10 meters into the ground. At this depth there should be seawater aquifers below the coast that regularly recharges itself with seawater from the ocean. A vertical well containing intake holes will be put into the borehole. The intake holes will be equipped with large size filters that can easily be replaced. At the surface, the well will be connected to a large-scale pump in a well pump house. From there, the seawater will be horizontally pumped into a seawater tank in the desalination plant. This will be considered the start of our process. The process ends after pretreatment, desalination, post treatment, and brine treatment are completed. Appendix 1 shows more details on the Design Basis.

The feed composition is assumed to be the standard seawater “average” based on the DOW Chemical ROSA 9.0 Technical Manual in Table 1. However, in reality the seawater composition at the plant site will have to be chemically tested. The final product will be based on L.A. standards for drinking water quality (Table 1). The specifications in the table are upper limits that we cannot exceed. We will aim to reach the salt level standards but any remineralization and pH treatment will be exported and paid for due to the added complexity and cost in making such a system in-house.

Due to the acidity of the desalinated water coming out of the reverse osmosis (RO) system (pH ~ 6), the pumps and pipes carrying the desalinated water for post treatment will likely corrode over time. We have decided to place the desalination plant by the Long Beach Water Reclamation Plant (about 50 meters away) to prevent pipe corrosion.

[[File:Table_1.jpg|center|200px|frame|Table 1. Concentration of the standard seawater, the output from the RO system, and LA standard.]]

=Desalination Overview and Alternatives=

We first began our research into the current landscape of the desalination industry, particularly of plants that were built in the United States. We discovered there were many processes but the main ones included: 
* Membrane Separation (Forward Osmosis and Reverse Osmosis)
* Thermal Separation (Multi Stage Flash Distillation)
* Phase Change Separation (Freeze Separation)
Each process was analyzed for their pros and cons. For our project design, we decided to go with reverse osmosis as our desalination process. This is because it is the most commonly used process in the industry, and for good reason. Unlike thermal separation, which requires huge amounts of energy to supply heat, reverse osmosis is very energy efficient. This reduces energy costs that would be spread across the plant’s lifetime.<ref name = "Republic">Republic of Mauritius NY. "Technical Aspects of Desalination Plant". Sustainable Sanitation and Water Management. </ref>

Forward osmosis is a newer technology being developed that also uses membranes to separate water and salt and may be more energy efficient overall than reverse osmosis. This is because forward osmosis doesn’t require a pump to pressurize seawater through the membranes, but rather leaves a higher salt concentrated solution (AKA draw solution) on the other side of the membrane for the seawater to be drawn into. Then a second step is required to regenerate the concentrated draw solution and produce purified water.<ref>Mallinson, Alissa. "Study Shows Forward Osmosis Desalination Not Energy Efficient." MIT News [Boston]. July 23, 2014.</ref> We decided to move away from forward osmosis because it is a process that is still being refined and developed for desalination and lacks the data, programs, and technologies that reverse osmosis users benefit from. Additionally, a recent study from MIT has claimed that due to the inherent nature of forward osmosis requiring that seawater be drawn to a higher concentrated salt solution before the water and salt is separated, it cannot possibly be more energy efficient than reverse osmosis.<ref>O'Hern, Sean C., Karnik, Rohit. Nanofiltration across Defect-Sealed Nanoporous Monolayer Graphene. American Chemical Society 15.5 (2014); 3254-260. </ref>

Other processes such as freeze separation were too restrictive in their feed and process demands and required a lot of capital. Ultimately, reverse osmosis is simply the best process that exists for desalination and has been ranked the best when performance is evaluated in terms of technical, environmental, and economic aspects. It is a process that is flexible in water quantity, operation start-up and shut-off, and site location.

Nevertheless there are disadvantages to reverse osmosis that must be considered. The feed seawater must not be too salty and too low quality. The pressure to pump the seawater is enormous. The construction time is long for large plants. The capital costs are also high when compared with other processes such as thermal separation. But all of these disadvantages are ultimately outweighed by the advantages reverse osmosis provides.<ref name = "Republic"/>

=Process Overview and Alternatives=

See Appendix 2 (Figure 1) for the Process Flow Diagram (PFD). The process itself is detailed below. The stream flow rates are in Appendix 2 (Table 1). We assumed 10 meters of pipe for every stream except the inlet and outlet streams of the entire PFD.

==Pretreatment==

After the 10 meter vertical well pumps seawater from the Long Beach coast, the feed is pumped 500 meters to a seawater tank that makes it easy to shut-down the system and reduce the large volume of flow. The level control keeps the seawater from overfilling the tank. From there, part of the seawater goes to brine treatment while the rest go to pretreatment. The purpose of pretreatment in the desalination process is to prevent equipment damage by removing particles and microorganisms from the water. A variety of pollutants -- including calcium carbonate, sulfates, silica, clay, bacteria, and other colloids and microorganisms -- can disrupt the reverse osmosis process. Media filtration uses fine grains such as sand or anthracite to filter the particles. We used a large filter to catch larger particle material (e.g. sand). However, simple filtration is not enough to prevent colloid fouling. An ultrafilter (i.e. microfiltration) is used to remove virtually all suspended matter, effectively preventing colloid fouling. The costs associated with ultrafiltration are not prohibitive in terms of the project’s budget.<ref name = "Desalination">Desalination Pre-treatment. Lenntech.</ref>

Reverse osmosis systems can also be disrupted by the growth of microbiological films of bacteria on the film and pipe walls. They are typically treated as colloids during pretreatment, and most organisms are removed from the water by ultrafiltration, but this does not remove the smallest of organic particles. DBNPA is the most effective biocide, but these aren’t often used and have to be removed from the water. Chlorination is an effective way to remove bacteria, but it can cause degradation of the membrane, and must be removed before the reverse osmosis occurs. In our process we will use UV irradiation to remove any organic material that passes through the microfilters.<ref name = "Desalination"/>

Scaling can reduce the flux across the membrane and can build up in pipe corners, causing backup. It is caused by inorganic salts -- specifically sulfates and calcium and magnesium carbonate -- present in the water. Scaling can be prevented and/or controlled with acids, antiscalants, or ion exchange softening. Antiscalants are good for removing sulfates, but not much else. Acidification will only be effective at removing CaCO3 from water but, the decreased pH may deteriorate the membrane and pipes, so it may require future research.<ref name = "Desalination"/> Ion exchange is very expensive and requires more capital to undergo, but is overall the more effective option.

The ion exchanger is a pretreatment system which acts as a method to fight membrane scaling in reverse osmosis. The ion exchanger works by using sodium carboxylate as a resin which will react with Chlorine and Magnesium ions, removing them from the feed. The sodium is substituted out and enters the feed stream, increasing the overall sodium level and TDS. The lack of magnesium and chlorine ions means that there is nothing to react with the carbonates to form scaling salts on the membranes.

==Desalination (Reverse Osmosis)==

Two pumps (P-105 & P-106) take the adjusted feed seawater up to 58 bar. However, before this pumping occurs, about 70% of the original stream gets separated and heads into a PX Pressure Exchanger. This separated stream later joins together with the pumped stream to enter the reverse osmosis system at the required pressure of 58 bar. The streams can combine together as they have equal pressure. It is the concentrate stream from the output of the reverse osmosis system that transfers its energy to the adjusted feed in the PX Pressure Exchanger.

[[File:Figure_3.jpg|left|300px|frame|Figure 3. Inside look of pressure exchanger]]
The pressure of the concentrate stream is about 57 bar and it is fed into an exchanger which works by using vessel that resembles a small pipe within a rotating wheel. The low pressure (LP) feed enters the vessel at 1 bar (see Figure 3). The feed turns and is pushed out into an outlet pipe by the high pressure (HP) concentrate stream at ~57 bar. The feed is now a HP stream via this exchanged energy. The now LP concentrate stream turns and is pushed out into an outlet pipe by the LP feed entering the vessel.

98% of the energy within the concentrate stream can be recovered via the PX Pressure Exchanger.<ref>Desalination Product and Applications Catalog 2013. “Advance Technologies for Seawater Reverse Osmosis and Brackish Water Applications”. Energy Recovery, Inc. 2012-2014.</ref >Thus, the now HP feed stream is not quite 58 bar and will need to be pumped by a small booster pump (P-107).

When designing a reverse osmosis system, we need to make decisions on what specific kind of reverse osmosis system will be implemented. When making the decision, we investigated the advantages of different types of reverse osmosis among the following options: batch vs. continuous process, number of stages, and number of elements.<ref>FILMTEC Reverse Osmosis Membranes. “Water & Process Solutions Technical Manual.” Dow Chemical Company.</ref> Our decision was made based on process optimization which is described later in this report. The end result of that optimization was to go with a single stage, continuous process with five elements in series.

=Post Treatment=

The permeate is sent to post treatment after the salt content is analyzed and controlled if too high or too low due to any unexpected errors. As stated above, the post treatment process will be done 50 meters from our plant by the Long Beach Water Reclamation Plant. The reason we aren’t performing post treatment “in-house” is due to the increased cost it would take for us to build and run the equipment for post treatment. We have the option to use the government’s water treatment facilities that already exist and our 15 million gallons per day output is a relatively small addition to the large Long Beach Water Reclamation Plant. We will assume any costs to post treatment will be pushed to the consumer in the price of purchasing the drinking water from the Long Beach Water Reclamation Plant.

Post treatment at any facility works to bring water to “taste standards”. The water that is the permeate of our process is already meeting L.A. drinking standards as shown above in Table 1 after our mass balance calculations around the reverse osmosis system (Appendix 3, Table 1). The post treatment increases the pH from 6.3 to 7.3 with the addition of base and then adds minerals like fluoride, calcium, and magnesium back in to improve taste.

==Brine Treatment and Byproduct Waste==

A natural consequence of a reverse osmosis procedure is that a lot of salt will be leftover from the membrane separation in a very concentrated form. This is called “brine” and will have to be taken care of. According to EPA requirements, the salt concentration of water entering the environment (via land or sea) cannot be higher than 40,000 ppm.8 The salt concentration of our brine is much higher (around 70,000 ppm). Thus, the brine must be removed or treated. We decided to treat the brine by diluting it as removing brine can be difficult and may require lots of land (e.g. evaporation ponds).

After the concentrate exchanges energy in the energy exchanger, it becomes a low pressure stream. It is pumped into a dilution tank and mixes with 18.2 million gallons per day of seawater from the inlet to become diluted brine which is pumped 600 meters out to sea.

The waste byproducts are the solid buildup on the filters and membranes. The filters and membranes also have a set lifetime and thus will also need to be removed. The waste is all non-hazardous and thus can be put in the municipal dump.

=Process Details and Equipment Sizing=

==Filters and UV Treatment [Pretreatment]==
The cost of our ultrafiltration system is estimated using data from the University of Tennessee at Knoxville.<ref>Qiang He and R. Bruce Robinson Membrane Filtration Systems. “Membrane Module Costs.” University of Tennessee at Knoxville. December 2014.</ref> It estimated the cost of the cost of the ultrafiltration $0.057 per 1000 gallons and the construction and capital costs are $0.39 per gallon per day. The cost of UV treatment is based off of units available from buyultraviolet.com for units that can treat about 450 gallons per minute.<ref>Atlantic Ultraviolet Corporation. “Megatron Water Disinfection System.”</ref> Due to the water having to be in relatively small pipes for treatment we do not expect the costs to scale very much, so we believe these units to be an accurate cost estimate.

==Ion Exchanger [Pretreatment] ==

The ion exchanger was sized using data from DOW. Our process will use 60 vessels with ion exchange beds of volume 5.73 m3. With a flow rate of 30 bed volumes/hour (10,312.2 m3/hr) the system of beds has a capacity of 125% of the feed. This allows for periodic regeneration of the resin, a weak acid cation resin (DOW MAC-3 Resin). This weak cation resin will remove cationic calcium and magnesium from the water to prevent scaling. The mass flow was calculated via a mass balance on the ion exchanger (Appendix 3).

==Reverse Osmosis System [Desalination]==

Each of the 2607 Reverse Osmosis SW30HRLE-440i [Seawater 30 High Rejection Low Energy 440 ft2 interlocking cap membrane] units have the same dimensions. A, the net length, is 1016 mm. B, the total length including the endcap, is 1029 mm. C, the diameter of the filter element, is 201mm. D, the diameter of the permeate pathway, is 29 mm. For a single stage, we have five of the elements in series so the overall length of a single stage is expected to be roughly 5145 mm. The membrane material is polyamide thin film composite. It will last three years.
Operating Conditions [ROSA 9.0]: 
* Temperature: 30oC 
* Pressure: 58 bar 
* Flow Rate: 52.3 Million gal per day 
* Recovery: 28.7%

[[File:Figure_4.png|center|400px|frame|Figure 4. Diagram of Dow FILMTEC SW30HRLE-440i]]
Cleaning Schedule: 
* Conservative Estimate: Once every three months (three days available for cleaning)<ref name = "Bates"/> 
* DOW FILMTEC Cleaning Conditions 
** The normalized permeate flow drops 10%
* The normalized salt passage increases 5 - 10%

==Seawater Holding Tank and Dilution Tank [Brine Treatment]==

The initial seawater holding tank works as a buffer to prevent surges and help perform shutdowns quickly and easily. We expect to run continuously for 350 days a year with the rest of the available time spread out and used for shutting down the plant for cleaning membranes, filters, and the fixing of unexpected errors. There is a big tradeoff between having a holding tank with high retention time and having a holding tank with low cost. With that in mind, we erred on the side of lowering capital cost and went with a retention time of one hour. The volume could be found by multiplying the retention time with the feed flow rate (Stream 1). We assumed that a good height to width ratio for a large water tank would be 4:1.

A mass balance was done around the dilution tank to find what amount of seawater would be necessary to dilute the brine to 45,000 ppm salinity (Appendix 3; Table 3). With this the feed flow rate into the tank could be found. It was found that it takes about 20 minutes of retention time for good dilution of the brine in seawater.8 We again assumed that a good height to width ratio for a large water tank would be 4:1.

We assumed a corrosion allowance of 4 mm wall thickness. We calculated wall thickness using vessel pressure, hoop stress, longitudinal stress, and weld factor. Each tank is made with SS316 to protect against corrosion.<ref>Towler G, Sinnott R. Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013.</ref>

<center>
'''Table 2.''' Tank sizes.
[[File:Table_2.jpg|center|600px]]
</center>

=Process Optimization=
<center> '''Table 3.'''Optimization Criteria
[[File:Table_3.jpg|center|400px]]
</center>
In order to optimize the reverse osmosis system, numerous simulations were run. Three representative cases are reported as shown above (Table 3). The objective of the optimization was to minimize the required water price to break even in 10 years. In order to achieve this goal, the balance between the capital cost from the number of the filter elements and utility cost from energy consumption was weighed. Before the energy recovery system was implemented, the general trend was that if we have more filter elements in series, overall recovery increases but recovery rate per filter decreases, and we need more filter elements. While increasing the elements in series increase the required number of filters, and thereby increases the capital cost, we require less energy to produce the same amount of water. After the energy recovery system is implemented, the correlation between the number of elements in series and the required energy becomes more complicated because high pressure concentrate is used to pressurize the feed. However, the trade-off between the utility cost and capital cost still exists. Of the three cases shown above, single stage, five elements RO system turned out to be the most optimal solution because single stage, single element system required higher utility cost and two stage system slightly did not meet the LA standard. Detailed information which was used for the optimization is available in Appendix 3 and in Appendix 6.

=Ethical, Environmental, and Societal Considerations=

There are a few key considerations we have followed to meet ethical, environmental, and societal guidelines and requirements: 
* Meet L.A. Drinking Water Standards
** Correct ion concentration
* Moved near water treatment plant
** Pipes will not corrode from low pH
* Working below pressure limits
* Water will be locally served and meet taste standards
** Via Post Treatment mineralization
* Follows EPA Environmental Salinity Requirements
* Reduced Energy Requirements
** PX Pressure Exchanger
** Optimized RO System
** Reverse Osmosis is itself Energy Saving

For further environmental considerations, see effluent summary in Appendix 4.

=Economic Analysis=

The pricing of the majority of the process was completed using the Aspen Economic Analyzer, however we had determine costs for the ultrafiltration, reverse osmosis membranes, and the ultraviolet treatment was completed separately. This led to a total ISBL of $56 MM, OSBL of $22.4 MM, Engineering Costs of $22.4 MM and a Contingency of $7.8 MM. We also chose a working Capital of $10 MM. This working capital is sufficient because we have no raw materials costs. We also had FCOP of $5.2 MM and VCOP of $3.5 MM.

We originally were planning on pricing our water at $700 per acre-foot, but after doing economic analysis we had a 10 year NPV of $-76.2 MM. After determining did a sensitivity analysis and determined that if our water was priced at $2104 per acre-foot we could have an NPV of $0 after 10 years and if we were able to get the same price of $2257 per acre-foot as the plant in Carlsbad we would have a 10 year NPV of $8.3 MM which increases to $41.1 MM at 20 years. For more detailed data about the economic analysis see Appendix 6.

=Conclusion and Recommendations=

Our mission was to create an economically viable process that will desalinate water for human consumption in the L.A. County. We have found after research and optimization that the best process to undergo desalination is a continuous single stage, five elements reverse osmosis system. Our process can meet the required L.A. drinking water standard and we have set up plans for post treatment at a local water reclamation site. While meeting the drinking water standards, we also met our ethical, environmental, and social standards.

Nevertheless, at a sale price of $700 per acre-foot, our proposed desalination plant in Los Angeles, California looks to be a very expensive project to undergo. Without any sort of government assistance to carry the heavy capital costs of such a project, we would suggest the plant not be built. However, after performing a sensitivity analysis at higher (and less conservative) price points we reached a breakeven point in 10 years with a sale price of $2104 per acre-foot of drinking water. Research has shown that the droughts striking California look only to get worse in the near future. With limited freshwater resources, there is no option other than to consider desalination. Therefore it is reasonable to expect we could get near the same price deal as local desalination plants like the Carlsbad Plant ($2257 per acre-foot) As we have met all design constraints for this project, we hope to convince policymakers to adopt our project at least $2104 per acre-foot water. If this deal is reached, we recommend that this project and design be funded.

However, there are ways to improve our design with future work. Right now all the pipes diameters are based on having a flowrate of 1.5 m/s. But in reality, we’d need to use standard piping sizes to get more accurate pricing and energy costs. Also, in reality we would be purchasing some of these vessels in bulk quantities (i.e. bulk filters) so the price could be lowered with negotiation and economies of scale. We also have pumps that are providing very little pressure for their high cost. We can remove these pumps and simply run at higher pressures earlier. We can potentially reduce early energy costs by building a “floating” desalination plant in the sea to reduce pumping distance from the well.

=Appendices=

==Appendix 1: Design Basis==
[[File:Appendix_1.png|thumb|border|center|middle|800*800|]]

==Appendix 2: Process Flow Diagram==
[[File:Appendix_2.png|center]]
==Appendix 3: Mass and Energy Balance Tables==
'''Table 1'''. Mass balance on 1 stage 5 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T1.png|center]]

'''Table 2'''. Mass balance on 1 stage 1 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T2.png|center]]

'''Table 3'''. Mass balance on 2 stage 5,5 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T3.png|center]]

'''Table 4'''. 1 stage 5 elements CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T4.png|center]]

'''Table 5'''. 1 stage 1 element CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T5.png|center]]

'''Table 6'''. 2 stages 5,5 elements CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T6.png|center]]

'''Table 7'''. Energy balance on 1 stage 5 elements RO system
[[File:Appendix_3_T7.png|center]]

==Appendix 4: Effluent Summary==
[[File:Appendix_4.png|center]]

==Appendix 5: Cost of Production==
[[File:Appendix_5.png|center]]

==Appendix 6: Economic Anlaysis==
===Case I ===
Water is $700; RO with 1 stage and 5 elements
[[File:Appendix_6_C1.png|center]]

===Case II===
Water is $1400; RO with 1 stage and 5 elements
[[File:Appendix_6_C2.png|center]]

===Case III: Carlsbad ===
Water is $2257; RO with 1 stage and 5 elements
[[File:Appendix_6_C3.png|center]]

===Case IV: Breakeven (NPV = 0 at 10 years)===
Water is $2104; RO with 1 stage and 5 elements
[[File:Appendix_6_C4.png|center]]

===Case V: Breakeven (NPV = 0 at 10 years)===
Water is $2112; RO with 2 stages and 5,5 elements
[[File:Appendix_6_C5.png|center]]

===Case VI: Breakeven (NPV = 0 at 10 years)===
Water is $2138 with 1 stage and 1 element
[[File:Appendix_6_C6.png|center]]

==Appendix 7: ROSA Simulation Outputs==

'''Table 1'''. System detail of 1 stage 5 elements RO system
[[File:Appendix_7_T1.png|center]]

'''Table 2'''. System detail of 1 stage 1 elements RO system
[[File:Appendix_7_T2.png|center]]

'''Table 3'''. System detail of 2 stage 5,5 elements system
[[File:Appendix_7_T3.png|center]]

'''Table 4'''. 1 stage 5 elements RO system concentrations and pH in each stream
[[File:Appendix_7_T4.png|center]]

'''Table 5'''. 1 stage 1 element RO system concentration and pH in each stream
[[File:Appendix_7_T5.png|center]]

'''Table 6'''. 2 stage 5,5 elements RO system concentrations and pH in each stream
[[File:Appendix_7_T6.png|center]]

=References=
<references/>

Desalination - Team E

2016-03-12T03:24:39Z

Bts511: /* Appendix 1: Design Basis */

Group E Corporation

Authors: Hassan Ali, Woo Soo Choe, Brett Sleyster, Jake Stolley

Instructors: Fengqi You, David Wegerer

March 11, 2016

=Executive Summary=

Outlined in this report is the project model report for the proposed desalination plant in Los Angeles, California which will provide drinking water to local citizens using reverse osmosis process. This location was chosen due to the region’s propensity for droughts in recent years. The fresh water resources in California have plummeted as water prices have risen 3-4% each year. This is only expected to continue. Other companies have worked deals with the Californian government to create desalination plants that provide a new fresh water drinking source. This suggests that there is a strong market here with potential growth.

The plant will have an output of 15 million gallons per day, which is about medium size relative to other plants. Thus the build time will be about one year. The plant will run for 30 years, with 350 days per year of use. Shutdown time will be used for cleaning and maintenance of equipment. The plant will be built on Long Beach, 500 meters from the coast nearby the Long Beach Water Reclamation Plant (LBWP). The input will withdraw 70.5 million gallons per day of seawater that will be used for reverse osmosis and dilution of brine.

Reverse osmosis was chosen as the desalination method for its low energy cost and flexibility. We designed a process that will pretreat the seawater to remove waste, colloids, bacteria, and membrane-damaging scalants. Once this seawater is treated is pumped to heavy pressures through a reverse osmosis system. After doing process optimization under various system parameter changes, we found a continuous, single stage, five element reverse osmosis process to be best. It produces a near 30% recovery. We reduce energy costs by including a pressure exchanger between our concentrate outlet and adjusted feed into the reverse osmosis system.

After performing mass and energy balances, we were able to size equipment and find our permeate composition. It turns out our process will meet L.A. standards for drinking water quality. However, post treatment will be done by the local LBWP to make the water taste better and be non-caustic to pipes. This will reduce “in-house” costs. Many environmental considerations were made, to meet our environmental standards. We meet EPA salinity requirements for our brine waste by diluting it. We also don’t use a lot of energy and thus reduce our fossil fuel consumption and CO2 generation. To be safe, we work well under max pressure limits for our equipment. We make sure we are ethical and meet all standards that our local county requires of us.

After an economic evaluation at a conservative well-water price of $700 per acre-foot, we realize that our revenue is way too low to cover our costs. Capital costs are especially large with regards to our large water tanks. We find that we reach a 10 year NPV breakpoint of $2104 per acre-foot after performing a sensitivity analysis. The Carlsbad Plant in San Diego worked out a deal for $2257 per acre-foot.

As we have met all design constraints for this project, we hope to convince policymakers to adopt our project at least $2104 per acre-foot water. If this deal is reached, we recommend that this project and design be funded. There is future work to be done to improve costs such as performing pump optimization. However, this work is minimal and won’t harm our bottom line if we reach our proposed price points.

__TOC__

=Introduction=

Due to a combination of various factors--including rising human population, climate change, increased energy consumption, and land erosion--the availability of fresh water as a resource for human consumption has been dwindling. This has had a negative impact on the world, but this impact has been especially felt in areas with natural shortages in fresh water. Desert regions, such as Saudi Arabia, have resorted to desalinating seawater at the coast and transporting up to their cities. This encompasses their entire freshwater system as very few natural water sources remain. In the United States there is a region which has shown increasing signs of following in the footsteps of Saudi Arabia and becoming fully reliant on desalination: Los Angeles, California.

=Location and Market Analysis=

In this project, Los Angeles was chosen as the location our plant, because the area has been subject to a severe drought for the last four years, as shown in Figure 1. The rapid decrease in the available water has caused permanent damages to the ground, decreased the amount of water in the reservoir from 45% to 25%, and reduced the amount of hydroelectric power produced in the state. Such deficiency in usable water has increased the water prices in Los Angeles by 6-7% annually for the last decade and is expected to rise by 3.4% annually for typical customers for the next five years. For customers who use more than 20000 gallons a year, the bill is expected to increase by 34% by 2021. This price increase is occurring because of both drought and maintenance. <ref>Pacific Institute: The California Drought. 2014.</ref>

[[File:Figure_1.jpg|right|200px|frame|Figure 1. California Drought Index.]]
As of March 2015 the Metropolitan Water District of Southern California was willing to pay as much as $700 per acre-foot of water to farmers in the Sacramento Valley, so we assume this will be the target price of the water produced by the this desalination plant.<ref>Hessel, Phil. “Dry Southern California Offers Northern Farmers Top Dollar for Water”. NBC News. March 18, 2015. </ref>

Currently, there is a plant in Carlsbad which supplies drinking water to the San Diego area. The plant can produce up to 50 million gallons per day and is the nation's largest water desalination plant. The water from this plant costs $2257 per acre-foot for the first 48000 acre-feet and about $2000 for any additional acre-foot. The county signed a contract with this plant to purchase the first 48000 acre-feet of water for the next 30 years.<ref>Carlsbad Desalination Plant. “Project Overview”. January 2016.</ref> Also, there is another desalination plant in Tampa that produces about 25 million gallons of water per day, and over 2000 other public and industrial plants with production capacity of over 300000 gallons per day. When conducting the economic analysis our plant, such numbers presented by other plants must be considered, so our project can yield positive profit.<ref>Leven, Rachel. “U.S. Desalination Industry Grows Since 2000; See as Essential to Meeting Supply Needs”. Bloomberg BNA. August 21, 2013.</ref>

The profitability of the project is naturally heavily dependent on the price of water in California. If the drought ends, water supply will increase and the water price will drop significantly, but if the drought continues, the price of water in California could rise even more rapidly and allow desalination plants to be more profitable.<ref>Southern California Public Radio. “Where is California water use decreasing?”. December 2015.</ref> It may be difficult to make this plant profitable without a guaranteed purchase or other subsidy from the local governments because the presence of pre-existing plants maintains the water price relatively low. However, external sources of profit such as guaranteed purchase or subsidy are expected because the ramifications of not having desalination plant or other sources of water may be significantly more expensive than the subsidizing a desalination plant assuming the drought continues.<ref>Weiser, Matt. “Could desalination solve California’s water problem?”. The Sacramento Bee. October 18, 2014.</ref>

=Design Basis and Technical Approach=

[[File:Figure_2.jpg|left|200px|frame|Figure 2. Vertical Well Feed]]
Considering the number of pre-existing desalination plant and the potential risk of the drought ending, our plant will be tentatively designed to produce 15 million gallons of drinkable water per day to homes and businesses in the Los Angeles county area. With a recovery rate of near 30%, and a requirement for diluting any brine, the input feed flowrate will be 70.5 million gallons per day. Following the similarly built Carlsbad plant, we expect the plant to last 30 years. The project building timeline is one year based on our relatively small plant size. We will run for 350 days a year while shutting down for cleaning and maintenance once every 3 months (based on DOW standards).<ref name = "Bates">Bates, Wayne. “Cleaning your RO”. Hydranautics. </ref> The plant will be built 500 meters from the coastline on Long Beach. A borehole will be drilled 10 meters into the ground. At this depth there should be seawater aquifers below the coast that regularly recharges itself with seawater from the ocean. A vertical well containing intake holes will be put into the borehole. The intake holes will be equipped with large size filters that can easily be replaced. At the surface, the well will be connected to a large-scale pump in a well pump house. From there, the seawater will be horizontally pumped into a seawater tank in the desalination plant. This will be considered the start of our process. The process ends after pretreatment, desalination, post treatment, and brine treatment are completed. Appendix 1 shows more details on the Design Basis.

The feed composition is assumed to be the standard seawater “average” based on the DOW Chemical ROSA 9.0 Technical Manual in Table 1. However, in reality the seawater composition at the plant site will have to be chemically tested. The final product will be based on L.A. standards for drinking water quality (Table 1). The specifications in the table are upper limits that we cannot exceed. We will aim to reach the salt level standards but any remineralization and pH treatment will be exported and paid for due to the added complexity and cost in making such a system in-house.

Due to the acidity of the desalinated water coming out of the reverse osmosis (RO) system (pH ~ 6), the pumps and pipes carrying the desalinated water for post treatment will likely corrode over time. We have decided to place the desalination plant by the Long Beach Water Reclamation Plant (about 50 meters away) to prevent pipe corrosion.

[[File:Table_1.jpg|center|200px|frame|Table 1. Concentration of the standard seawater, the output from the RO system, and LA standard.]]

=Desalination Overview and Alternatives=

We first began our research into the current landscape of the desalination industry, particularly of plants that were built in the United States. We discovered there were many processes but the main ones included: 
* Membrane Separation (Forward Osmosis and Reverse Osmosis)
* Thermal Separation (Multi Stage Flash Distillation)
* Phase Change Separation (Freeze Separation)
Each process was analyzed for their pros and cons. For our project design, we decided to go with reverse osmosis as our desalination process. This is because it is the most commonly used process in the industry, and for good reason. Unlike thermal separation, which requires huge amounts of energy to supply heat, reverse osmosis is very energy efficient. This reduces energy costs that would be spread across the plant’s lifetime.<ref name = "Republic">Republic of Mauritius NY. "Technical Aspects of Desalination Plant". Sustainable Sanitation and Water Management. </ref>

Forward osmosis is a newer technology being developed that also uses membranes to separate water and salt and may be more energy efficient overall than reverse osmosis. This is because forward osmosis doesn’t require a pump to pressurize seawater through the membranes, but rather leaves a higher salt concentrated solution (AKA draw solution) on the other side of the membrane for the seawater to be drawn into. Then a second step is required to regenerate the concentrated draw solution and produce purified water.<ref>Mallinson, Alissa. "Study Shows Forward Osmosis Desalination Not Energy Efficient." MIT News [Boston]. July 23, 2014.</ref> We decided to move away from forward osmosis because it is a process that is still being refined and developed for desalination and lacks the data, programs, and technologies that reverse osmosis users benefit from. Additionally, a recent study from MIT has claimed that due to the inherent nature of forward osmosis requiring that seawater be drawn to a higher concentrated salt solution before the water and salt is separated, it cannot possibly be more energy efficient than reverse osmosis.<ref>O'Hern, Sean C., Karnik, Rohit. Nanofiltration across Defect-Sealed Nanoporous Monolayer Graphene. American Chemical Society 15.5 (2014); 3254-260. </ref>

Other processes such as freeze separation were too restrictive in their feed and process demands and required a lot of capital. Ultimately, reverse osmosis is simply the best process that exists for desalination and has been ranked the best when performance is evaluated in terms of technical, environmental, and economic aspects. It is a process that is flexible in water quantity, operation start-up and shut-off, and site location.

Nevertheless there are disadvantages to reverse osmosis that must be considered. The feed seawater must not be too salty and too low quality. The pressure to pump the seawater is enormous. The construction time is long for large plants. The capital costs are also high when compared with other processes such as thermal separation. But all of these disadvantages are ultimately outweighed by the advantages reverse osmosis provides.<ref name = "Republic"/>

=Process Overview and Alternatives=

See Appendix 2 (Figure 1) for the Process Flow Diagram (PFD). The process itself is detailed below. The stream flow rates are in Appendix 2 (Table 1). We assumed 10 meters of pipe for every stream except the inlet and outlet streams of the entire PFD.

==Pretreatment==

After the 10 meter vertical well pumps seawater from the Long Beach coast, the feed is pumped 500 meters to a seawater tank that makes it easy to shut-down the system and reduce the large volume of flow. The level control keeps the seawater from overfilling the tank. From there, part of the seawater goes to brine treatment while the rest go to pretreatment. The purpose of pretreatment in the desalination process is to prevent equipment damage by removing particles and microorganisms from the water. A variety of pollutants -- including calcium carbonate, sulfates, silica, clay, bacteria, and other colloids and microorganisms -- can disrupt the reverse osmosis process. Media filtration uses fine grains such as sand or anthracite to filter the particles. We used a large filter to catch larger particle material (e.g. sand). However, simple filtration is not enough to prevent colloid fouling. An ultrafilter (i.e. microfiltration) is used to remove virtually all suspended matter, effectively preventing colloid fouling. The costs associated with ultrafiltration are not prohibitive in terms of the project’s budget.<ref name = "Desalination">Desalination Pre-treatment. Lenntech.</ref>

Reverse osmosis systems can also be disrupted by the growth of microbiological films of bacteria on the film and pipe walls. They are typically treated as colloids during pretreatment, and most organisms are removed from the water by ultrafiltration, but this does not remove the smallest of organic particles. DBNPA is the most effective biocide, but these aren’t often used and have to be removed from the water. Chlorination is an effective way to remove bacteria, but it can cause degradation of the membrane, and must be removed before the reverse osmosis occurs. In our process we will use UV irradiation to remove any organic material that passes through the microfilters.<ref name = "Desalination"/>

Scaling can reduce the flux across the membrane and can build up in pipe corners, causing backup. It is caused by inorganic salts -- specifically sulfates and calcium and magnesium carbonate -- present in the water. Scaling can be prevented and/or controlled with acids, antiscalants, or ion exchange softening. Antiscalants are good for removing sulfates, but not much else. Acidification will only be effective at removing CaCO3 from water but, the decreased pH may deteriorate the membrane and pipes, so it may require future research.<ref name = "Desalination"/> Ion exchange is very expensive and requires more capital to undergo, but is overall the more effective option.

The ion exchanger is a pretreatment system which acts as a method to fight membrane scaling in reverse osmosis. The ion exchanger works by using sodium carboxylate as a resin which will react with Chlorine and Magnesium ions, removing them from the feed. The sodium is substituted out and enters the feed stream, increasing the overall sodium level and TDS. The lack of magnesium and chlorine ions means that there is nothing to react with the carbonates to form scaling salts on the membranes.

==Desalination (Reverse Osmosis)==

Two pumps (P-105 & P-106) take the adjusted feed seawater up to 58 bar. However, before this pumping occurs, about 70% of the original stream gets separated and heads into a PX Pressure Exchanger. This separated stream later joins together with the pumped stream to enter the reverse osmosis system at the required pressure of 58 bar. The streams can combine together as they have equal pressure. It is the concentrate stream from the output of the reverse osmosis system that transfers its energy to the adjusted feed in the PX Pressure Exchanger.

[[File:Figure_3.jpg|left|300px|frame|Figure 3. Inside look of pressure exchanger]]
The pressure of the concentrate stream is about 57 bar and it is fed into an exchanger which works by using vessel that resembles a small pipe within a rotating wheel. The low pressure (LP) feed enters the vessel at 1 bar (see Figure 3). The feed turns and is pushed out into an outlet pipe by the high pressure (HP) concentrate stream at ~57 bar. The feed is now a HP stream via this exchanged energy. The now LP concentrate stream turns and is pushed out into an outlet pipe by the LP feed entering the vessel.

98% of the energy within the concentrate stream can be recovered via the PX Pressure Exchanger.<ref>Desalination Product and Applications Catalog 2013. “Advance Technologies for Seawater Reverse Osmosis and Brackish Water Applications”. Energy Recovery, Inc. 2012-2014.</ref >Thus, the now HP feed stream is not quite 58 bar and will need to be pumped by a small booster pump (P-107).

When designing a reverse osmosis system, we need to make decisions on what specific kind of reverse osmosis system will be implemented. When making the decision, we investigated the advantages of different types of reverse osmosis among the following options: batch vs. continuous process, number of stages, and number of elements.<ref>FILMTEC Reverse Osmosis Membranes. “Water & Process Solutions Technical Manual.” Dow Chemical Company.</ref> Our decision was made based on process optimization which is described later in this report. The end result of that optimization was to go with a single stage, continuous process with five elements in series.

=Post Treatment=

The permeate is sent to post treatment after the salt content is analyzed and controlled if too high or too low due to any unexpected errors. As stated above, the post treatment process will be done 50 meters from our plant by the Long Beach Water Reclamation Plant. The reason we aren’t performing post treatment “in-house” is due to the increased cost it would take for us to build and run the equipment for post treatment. We have the option to use the government’s water treatment facilities that already exist and our 15 million gallons per day output is a relatively small addition to the large Long Beach Water Reclamation Plant. We will assume any costs to post treatment will be pushed to the consumer in the price of purchasing the drinking water from the Long Beach Water Reclamation Plant.

Post treatment at any facility works to bring water to “taste standards”. The water that is the permeate of our process is already meeting L.A. drinking standards as shown above in Table 1 after our mass balance calculations around the reverse osmosis system (Appendix 3, Table 1). The post treatment increases the pH from 6.3 to 7.3 with the addition of base and then adds minerals like fluoride, calcium, and magnesium back in to improve taste.

==Brine Treatment and Byproduct Waste==

A natural consequence of a reverse osmosis procedure is that a lot of salt will be leftover from the membrane separation in a very concentrated form. This is called “brine” and will have to be taken care of. According to EPA requirements, the salt concentration of water entering the environment (via land or sea) cannot be higher than 40,000 ppm.8 The salt concentration of our brine is much higher (around 70,000 ppm). Thus, the brine must be removed or treated. We decided to treat the brine by diluting it as removing brine can be difficult and may require lots of land (e.g. evaporation ponds).

After the concentrate exchanges energy in the energy exchanger, it becomes a low pressure stream. It is pumped into a dilution tank and mixes with 18.2 million gallons per day of seawater from the inlet to become diluted brine which is pumped 600 meters out to sea.

The waste byproducts are the solid buildup on the filters and membranes. The filters and membranes also have a set lifetime and thus will also need to be removed. The waste is all non-hazardous and thus can be put in the municipal dump.

=Process Details and Equipment Sizing=

==Filters and UV Treatment [Pretreatment]==
The cost of our ultrafiltration system is estimated using data from the University of Tennessee at Knoxville.<ref>Qiang He and R. Bruce Robinson Membrane Filtration Systems. “Membrane Module Costs.” University of Tennessee at Knoxville. December 2014.</ref> It estimated the cost of the cost of the ultrafiltration $0.057 per 1000 gallons and the construction and capital costs are $0.39 per gallon per day. The cost of UV treatment is based off of units available from buyultraviolet.com for units that can treat about 450 gallons per minute.<ref>Atlantic Ultraviolet Corporation. “Megatron Water Disinfection System.”</ref> Due to the water having to be in relatively small pipes for treatment we do not expect the costs to scale very much, so we believe these units to be an accurate cost estimate.

==Ion Exchanger [Pretreatment] ==

The ion exchanger was sized using data from DOW. Our process will use 60 vessels with ion exchange beds of volume 5.73 m3. With a flow rate of 30 bed volumes/hour (10,312.2 m3/hr) the system of beds has a capacity of 125% of the feed. This allows for periodic regeneration of the resin, a weak acid cation resin (DOW MAC-3 Resin). This weak cation resin will remove cationic calcium and magnesium from the water to prevent scaling. The mass flow was calculated via a mass balance on the ion exchanger (Appendix 3).

==Reverse Osmosis System [Desalination]==

Each of the 2607 Reverse Osmosis SW30HRLE-440i [Seawater 30 High Rejection Low Energy 440 ft2 interlocking cap membrane] units have the same dimensions. A, the net length, is 1016 mm. B, the total length including the endcap, is 1029 mm. C, the diameter of the filter element, is 201mm. D, the diameter of the permeate pathway, is 29 mm. For a single stage, we have five of the elements in series so the overall length of a single stage is expected to be roughly 5145 mm. The membrane material is polyamide thin film composite. It will last three years.
Operating Conditions [ROSA 9.0]: 
* Temperature: 30oC 
* Pressure: 58 bar 
* Flow Rate: 52.3 Million gal per day 
* Recovery: 28.7%

[[File:Figure_4.png|center|400px|frame|Figure 4. Diagram of Dow FILMTEC SW30HRLE-440i]]
Cleaning Schedule: 
* Conservative Estimate: Once every three months (three days available for cleaning)<ref name = "Bates"/> 
* DOW FILMTEC Cleaning Conditions 
** The normalized permeate flow drops 10%
* The normalized salt passage increases 5 - 10%

==Seawater Holding Tank and Dilution Tank [Brine Treatment]==

The initial seawater holding tank works as a buffer to prevent surges and help perform shutdowns quickly and easily. We expect to run continuously for 350 days a year with the rest of the available time spread out and used for shutting down the plant for cleaning membranes, filters, and the fixing of unexpected errors. There is a big tradeoff between having a holding tank with high retention time and having a holding tank with low cost. With that in mind, we erred on the side of lowering capital cost and went with a retention time of one hour. The volume could be found by multiplying the retention time with the feed flow rate (Stream 1). We assumed that a good height to width ratio for a large water tank would be 4:1.

A mass balance was done around the dilution tank to find what amount of seawater would be necessary to dilute the brine to 45,000 ppm salinity (Appendix 3; Table 3). With this the feed flow rate into the tank could be found. It was found that it takes about 20 minutes of retention time for good dilution of the brine in seawater.8 We again assumed that a good height to width ratio for a large water tank would be 4:1.

We assumed a corrosion allowance of 4 mm wall thickness. We calculated wall thickness using vessel pressure, hoop stress, longitudinal stress, and weld factor. Each tank is made with SS316 to protect against corrosion.<ref>Towler G, Sinnott R. Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013.</ref>

<center>
'''Table 2.''' Tank sizes.
[[File:Table_2.jpg|center|600px]]
</center>

=Process Optimization=
<center> '''Table 3.'''Optimization Criteria
[[File:Table_3.jpg|center|400px]]
</center>
In order to optimize the reverse osmosis system, numerous simulations were run. Three representative cases are reported as shown above (Table 3). The objective of the optimization was to minimize the required water price to break even in 10 years. In order to achieve this goal, the balance between the capital cost from the number of the filter elements and utility cost from energy consumption was weighed. Before the energy recovery system was implemented, the general trend was that if we have more filter elements in series, overall recovery increases but recovery rate per filter decreases, and we need more filter elements. While increasing the elements in series increase the required number of filters, and thereby increases the capital cost, we require less energy to produce the same amount of water. After the energy recovery system is implemented, the correlation between the number of elements in series and the required energy becomes more complicated because high pressure concentrate is used to pressurize the feed. However, the trade-off between the utility cost and capital cost still exists. Of the three cases shown above, single stage, five elements RO system turned out to be the most optimal solution because single stage, single element system required higher utility cost and two stage system slightly did not meet the LA standard. Detailed information which was used for the optimization is available in Appendix 3 and in Appendix 6.

=Ethical, Environmental, and Societal Considerations=

There are a few key considerations we have followed to meet ethical, environmental, and societal guidelines and requirements: 
* Meet L.A. Drinking Water Standards
** Correct ion concentration
* Moved near water treatment plant
** Pipes will not corrode from low pH
* Working below pressure limits
* Water will be locally served and meet taste standards
** Via Post Treatment mineralization
* Follows EPA Environmental Salinity Requirements
* Reduced Energy Requirements
** PX Pressure Exchanger
** Optimized RO System
** Reverse Osmosis is itself Energy Saving

For further environmental considerations, see effluent summary in Appendix 4.

=Economic Analysis=

The pricing of the majority of the process was completed using the Aspen Economic Analyzer, however we had determine costs for the ultrafiltration, reverse osmosis membranes, and the ultraviolet treatment was completed separately. This led to a total ISBL of $56 MM, OSBL of $22.4 MM, Engineering Costs of $22.4 MM and a Contingency of $7.8 MM. We also chose a working Capital of $10 MM. This working capital is sufficient because we have no raw materials costs. We also had FCOP of $5.2 MM and VCOP of $3.5 MM.

We originally were planning on pricing our water at $700 per acre-foot, but after doing economic analysis we had a 10 year NPV of $-76.2 MM. After determining did a sensitivity analysis and determined that if our water was priced at $2104 per acre-foot we could have an NPV of $0 after 10 years and if we were able to get the same price of $2257 per acre-foot as the plant in Carlsbad we would have a 10 year NPV of $8.3 MM which increases to $41.1 MM at 20 years. For more detailed data about the economic analysis see Appendix 6.

=Conclusion and Recommendations=

Our mission was to create an economically viable process that will desalinate water for human consumption in the L.A. County. We have found after research and optimization that the best process to undergo desalination is a continuous single stage, five elements reverse osmosis system. Our process can meet the required L.A. drinking water standard and we have set up plans for post treatment at a local water reclamation site. While meeting the drinking water standards, we also met our ethical, environmental, and social standards.

Nevertheless, at a sale price of $700 per acre-foot, our proposed desalination plant in Los Angeles, California looks to be a very expensive project to undergo. Without any sort of government assistance to carry the heavy capital costs of such a project, we would suggest the plant not be built. However, after performing a sensitivity analysis at higher (and less conservative) price points we reached a breakeven point in 10 years with a sale price of $2104 per acre-foot of drinking water. Research has shown that the droughts striking California look only to get worse in the near future. With limited freshwater resources, there is no option other than to consider desalination. Therefore it is reasonable to expect we could get near the same price deal as local desalination plants like the Carlsbad Plant ($2257 per acre-foot) As we have met all design constraints for this project, we hope to convince policymakers to adopt our project at least $2104 per acre-foot water. If this deal is reached, we recommend that this project and design be funded.

However, there are ways to improve our design with future work. Right now all the pipes diameters are based on having a flowrate of 1.5 m/s. But in reality, we’d need to use standard piping sizes to get more accurate pricing and energy costs. Also, in reality we would be purchasing some of these vessels in bulk quantities (i.e. bulk filters) so the price could be lowered with negotiation and economies of scale. We also have pumps that are providing very little pressure for their high cost. We can remove these pumps and simply run at higher pressures earlier. We can potentially reduce early energy costs by building a “floating” desalination plant in the sea to reduce pumping distance from the well.

=Appendices=

==Appendix 1: Design Basis==
[[File:Appendix_1.png|thumb|border|center|middle|300*500|]]

==Appendix 2: Process Flow Diagram==
[[File:Appendix_2.png|center]]
==Appendix 3: Mass and Energy Balance Tables==
'''Table 1'''. Mass balance on 1 stage 5 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T1.png|center]]

'''Table 2'''. Mass balance on 1 stage 1 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T2.png|center]]

'''Table 3'''. Mass balance on 2 stage 5,5 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T3.png|center]]

'''Table 4'''. 1 stage 5 elements CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T4.png|center]]

'''Table 5'''. 1 stage 1 element CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T5.png|center]]

'''Table 6'''. 2 stages 5,5 elements CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T6.png|center]]

'''Table 7'''. Energy balance on 1 stage 5 elements RO system
[[File:Appendix_3_T7.png|center]]

==Appendix 4: Effluent Summary==
[[File:Appendix_4.png|center]]

==Appendix 5: Cost of Production==
[[File:Appendix_5.png|center]]

==Appendix 6: Economic Anlaysis==
===Case I ===
Water is $700; RO with 1 stage and 5 elements
[[File:Appendix_6_C1.png|center]]

===Case II===
Water is $1400; RO with 1 stage and 5 elements
[[File:Appendix_6_C2.png|center]]

===Case III: Carlsbad ===
Water is $2257; RO with 1 stage and 5 elements
[[File:Appendix_6_C3.png|center]]

===Case IV: Breakeven (NPV = 0 at 10 years)===
Water is $2104; RO with 1 stage and 5 elements
[[File:Appendix_6_C4.png|center]]

===Case V: Breakeven (NPV = 0 at 10 years)===
Water is $2112; RO with 2 stages and 5,5 elements
[[File:Appendix_6_C5.png|center]]

===Case VI: Breakeven (NPV = 0 at 10 years)===
Water is $2138 with 1 stage and 1 element
[[File:Appendix_6_C6.png|center]]

==Appendix 7: ROSA Simulation Outputs==

'''Table 1'''. System detail of 1 stage 5 elements RO system
[[File:Appendix_7_T1.png|center]]

'''Table 2'''. System detail of 1 stage 1 elements RO system
[[File:Appendix_7_T2.png|center]]

'''Table 3'''. System detail of 2 stage 5,5 elements system
[[File:Appendix_7_T3.png|center]]

'''Table 4'''. 1 stage 5 elements RO system concentrations and pH in each stream
[[File:Appendix_7_T4.png|center]]

'''Table 5'''. 1 stage 1 element RO system concentration and pH in each stream
[[File:Appendix_7_T5.png|center]]

'''Table 6'''. 2 stage 5,5 elements RO system concentrations and pH in each stream
[[File:Appendix_7_T6.png|center]]

=References=
<references/>

Desalination - Team E

2016-03-12T03:24:18Z

Bts511: /* Appendix 1: Design Basis */

Group E Corporation

Authors: Hassan Ali, Woo Soo Choe, Brett Sleyster, Jake Stolley

Instructors: Fengqi You, David Wegerer

March 11, 2016

=Executive Summary=

Outlined in this report is the project model report for the proposed desalination plant in Los Angeles, California which will provide drinking water to local citizens using reverse osmosis process. This location was chosen due to the region’s propensity for droughts in recent years. The fresh water resources in California have plummeted as water prices have risen 3-4% each year. This is only expected to continue. Other companies have worked deals with the Californian government to create desalination plants that provide a new fresh water drinking source. This suggests that there is a strong market here with potential growth.

The plant will have an output of 15 million gallons per day, which is about medium size relative to other plants. Thus the build time will be about one year. The plant will run for 30 years, with 350 days per year of use. Shutdown time will be used for cleaning and maintenance of equipment. The plant will be built on Long Beach, 500 meters from the coast nearby the Long Beach Water Reclamation Plant (LBWP). The input will withdraw 70.5 million gallons per day of seawater that will be used for reverse osmosis and dilution of brine.

Reverse osmosis was chosen as the desalination method for its low energy cost and flexibility. We designed a process that will pretreat the seawater to remove waste, colloids, bacteria, and membrane-damaging scalants. Once this seawater is treated is pumped to heavy pressures through a reverse osmosis system. After doing process optimization under various system parameter changes, we found a continuous, single stage, five element reverse osmosis process to be best. It produces a near 30% recovery. We reduce energy costs by including a pressure exchanger between our concentrate outlet and adjusted feed into the reverse osmosis system.

After performing mass and energy balances, we were able to size equipment and find our permeate composition. It turns out our process will meet L.A. standards for drinking water quality. However, post treatment will be done by the local LBWP to make the water taste better and be non-caustic to pipes. This will reduce “in-house” costs. Many environmental considerations were made, to meet our environmental standards. We meet EPA salinity requirements for our brine waste by diluting it. We also don’t use a lot of energy and thus reduce our fossil fuel consumption and CO2 generation. To be safe, we work well under max pressure limits for our equipment. We make sure we are ethical and meet all standards that our local county requires of us.

After an economic evaluation at a conservative well-water price of $700 per acre-foot, we realize that our revenue is way too low to cover our costs. Capital costs are especially large with regards to our large water tanks. We find that we reach a 10 year NPV breakpoint of $2104 per acre-foot after performing a sensitivity analysis. The Carlsbad Plant in San Diego worked out a deal for $2257 per acre-foot.

As we have met all design constraints for this project, we hope to convince policymakers to adopt our project at least $2104 per acre-foot water. If this deal is reached, we recommend that this project and design be funded. There is future work to be done to improve costs such as performing pump optimization. However, this work is minimal and won’t harm our bottom line if we reach our proposed price points.

__TOC__

=Introduction=

Due to a combination of various factors--including rising human population, climate change, increased energy consumption, and land erosion--the availability of fresh water as a resource for human consumption has been dwindling. This has had a negative impact on the world, but this impact has been especially felt in areas with natural shortages in fresh water. Desert regions, such as Saudi Arabia, have resorted to desalinating seawater at the coast and transporting up to their cities. This encompasses their entire freshwater system as very few natural water sources remain. In the United States there is a region which has shown increasing signs of following in the footsteps of Saudi Arabia and becoming fully reliant on desalination: Los Angeles, California.

=Location and Market Analysis=

In this project, Los Angeles was chosen as the location our plant, because the area has been subject to a severe drought for the last four years, as shown in Figure 1. The rapid decrease in the available water has caused permanent damages to the ground, decreased the amount of water in the reservoir from 45% to 25%, and reduced the amount of hydroelectric power produced in the state. Such deficiency in usable water has increased the water prices in Los Angeles by 6-7% annually for the last decade and is expected to rise by 3.4% annually for typical customers for the next five years. For customers who use more than 20000 gallons a year, the bill is expected to increase by 34% by 2021. This price increase is occurring because of both drought and maintenance. <ref>Pacific Institute: The California Drought. 2014.</ref>

[[File:Figure_1.jpg|right|200px|frame|Figure 1. California Drought Index.]]
As of March 2015 the Metropolitan Water District of Southern California was willing to pay as much as $700 per acre-foot of water to farmers in the Sacramento Valley, so we assume this will be the target price of the water produced by the this desalination plant.<ref>Hessel, Phil. “Dry Southern California Offers Northern Farmers Top Dollar for Water”. NBC News. March 18, 2015. </ref>

Currently, there is a plant in Carlsbad which supplies drinking water to the San Diego area. The plant can produce up to 50 million gallons per day and is the nation's largest water desalination plant. The water from this plant costs $2257 per acre-foot for the first 48000 acre-feet and about $2000 for any additional acre-foot. The county signed a contract with this plant to purchase the first 48000 acre-feet of water for the next 30 years.<ref>Carlsbad Desalination Plant. “Project Overview”. January 2016.</ref> Also, there is another desalination plant in Tampa that produces about 25 million gallons of water per day, and over 2000 other public and industrial plants with production capacity of over 300000 gallons per day. When conducting the economic analysis our plant, such numbers presented by other plants must be considered, so our project can yield positive profit.<ref>Leven, Rachel. “U.S. Desalination Industry Grows Since 2000; See as Essential to Meeting Supply Needs”. Bloomberg BNA. August 21, 2013.</ref>

The profitability of the project is naturally heavily dependent on the price of water in California. If the drought ends, water supply will increase and the water price will drop significantly, but if the drought continues, the price of water in California could rise even more rapidly and allow desalination plants to be more profitable.<ref>Southern California Public Radio. “Where is California water use decreasing?”. December 2015.</ref> It may be difficult to make this plant profitable without a guaranteed purchase or other subsidy from the local governments because the presence of pre-existing plants maintains the water price relatively low. However, external sources of profit such as guaranteed purchase or subsidy are expected because the ramifications of not having desalination plant or other sources of water may be significantly more expensive than the subsidizing a desalination plant assuming the drought continues.<ref>Weiser, Matt. “Could desalination solve California’s water problem?”. The Sacramento Bee. October 18, 2014.</ref>

=Design Basis and Technical Approach=

[[File:Figure_2.jpg|left|200px|frame|Figure 2. Vertical Well Feed]]
Considering the number of pre-existing desalination plant and the potential risk of the drought ending, our plant will be tentatively designed to produce 15 million gallons of drinkable water per day to homes and businesses in the Los Angeles county area. With a recovery rate of near 30%, and a requirement for diluting any brine, the input feed flowrate will be 70.5 million gallons per day. Following the similarly built Carlsbad plant, we expect the plant to last 30 years. The project building timeline is one year based on our relatively small plant size. We will run for 350 days a year while shutting down for cleaning and maintenance once every 3 months (based on DOW standards).<ref name = "Bates">Bates, Wayne. “Cleaning your RO”. Hydranautics. </ref> The plant will be built 500 meters from the coastline on Long Beach. A borehole will be drilled 10 meters into the ground. At this depth there should be seawater aquifers below the coast that regularly recharges itself with seawater from the ocean. A vertical well containing intake holes will be put into the borehole. The intake holes will be equipped with large size filters that can easily be replaced. At the surface, the well will be connected to a large-scale pump in a well pump house. From there, the seawater will be horizontally pumped into a seawater tank in the desalination plant. This will be considered the start of our process. The process ends after pretreatment, desalination, post treatment, and brine treatment are completed. Appendix 1 shows more details on the Design Basis.

The feed composition is assumed to be the standard seawater “average” based on the DOW Chemical ROSA 9.0 Technical Manual in Table 1. However, in reality the seawater composition at the plant site will have to be chemically tested. The final product will be based on L.A. standards for drinking water quality (Table 1). The specifications in the table are upper limits that we cannot exceed. We will aim to reach the salt level standards but any remineralization and pH treatment will be exported and paid for due to the added complexity and cost in making such a system in-house.

Due to the acidity of the desalinated water coming out of the reverse osmosis (RO) system (pH ~ 6), the pumps and pipes carrying the desalinated water for post treatment will likely corrode over time. We have decided to place the desalination plant by the Long Beach Water Reclamation Plant (about 50 meters away) to prevent pipe corrosion.

[[File:Table_1.jpg|center|200px|frame|Table 1. Concentration of the standard seawater, the output from the RO system, and LA standard.]]

=Desalination Overview and Alternatives=

We first began our research into the current landscape of the desalination industry, particularly of plants that were built in the United States. We discovered there were many processes but the main ones included: 
* Membrane Separation (Forward Osmosis and Reverse Osmosis)
* Thermal Separation (Multi Stage Flash Distillation)
* Phase Change Separation (Freeze Separation)
Each process was analyzed for their pros and cons. For our project design, we decided to go with reverse osmosis as our desalination process. This is because it is the most commonly used process in the industry, and for good reason. Unlike thermal separation, which requires huge amounts of energy to supply heat, reverse osmosis is very energy efficient. This reduces energy costs that would be spread across the plant’s lifetime.<ref name = "Republic">Republic of Mauritius NY. "Technical Aspects of Desalination Plant". Sustainable Sanitation and Water Management. </ref>

Forward osmosis is a newer technology being developed that also uses membranes to separate water and salt and may be more energy efficient overall than reverse osmosis. This is because forward osmosis doesn’t require a pump to pressurize seawater through the membranes, but rather leaves a higher salt concentrated solution (AKA draw solution) on the other side of the membrane for the seawater to be drawn into. Then a second step is required to regenerate the concentrated draw solution and produce purified water.<ref>Mallinson, Alissa. "Study Shows Forward Osmosis Desalination Not Energy Efficient." MIT News [Boston]. July 23, 2014.</ref> We decided to move away from forward osmosis because it is a process that is still being refined and developed for desalination and lacks the data, programs, and technologies that reverse osmosis users benefit from. Additionally, a recent study from MIT has claimed that due to the inherent nature of forward osmosis requiring that seawater be drawn to a higher concentrated salt solution before the water and salt is separated, it cannot possibly be more energy efficient than reverse osmosis.<ref>O'Hern, Sean C., Karnik, Rohit. Nanofiltration across Defect-Sealed Nanoporous Monolayer Graphene. American Chemical Society 15.5 (2014); 3254-260. </ref>

Other processes such as freeze separation were too restrictive in their feed and process demands and required a lot of capital. Ultimately, reverse osmosis is simply the best process that exists for desalination and has been ranked the best when performance is evaluated in terms of technical, environmental, and economic aspects. It is a process that is flexible in water quantity, operation start-up and shut-off, and site location.

Nevertheless there are disadvantages to reverse osmosis that must be considered. The feed seawater must not be too salty and too low quality. The pressure to pump the seawater is enormous. The construction time is long for large plants. The capital costs are also high when compared with other processes such as thermal separation. But all of these disadvantages are ultimately outweighed by the advantages reverse osmosis provides.<ref name = "Republic"/>

=Process Overview and Alternatives=

See Appendix 2 (Figure 1) for the Process Flow Diagram (PFD). The process itself is detailed below. The stream flow rates are in Appendix 2 (Table 1). We assumed 10 meters of pipe for every stream except the inlet and outlet streams of the entire PFD.

==Pretreatment==

After the 10 meter vertical well pumps seawater from the Long Beach coast, the feed is pumped 500 meters to a seawater tank that makes it easy to shut-down the system and reduce the large volume of flow. The level control keeps the seawater from overfilling the tank. From there, part of the seawater goes to brine treatment while the rest go to pretreatment. The purpose of pretreatment in the desalination process is to prevent equipment damage by removing particles and microorganisms from the water. A variety of pollutants -- including calcium carbonate, sulfates, silica, clay, bacteria, and other colloids and microorganisms -- can disrupt the reverse osmosis process. Media filtration uses fine grains such as sand or anthracite to filter the particles. We used a large filter to catch larger particle material (e.g. sand). However, simple filtration is not enough to prevent colloid fouling. An ultrafilter (i.e. microfiltration) is used to remove virtually all suspended matter, effectively preventing colloid fouling. The costs associated with ultrafiltration are not prohibitive in terms of the project’s budget.<ref name = "Desalination">Desalination Pre-treatment. Lenntech.</ref>

Reverse osmosis systems can also be disrupted by the growth of microbiological films of bacteria on the film and pipe walls. They are typically treated as colloids during pretreatment, and most organisms are removed from the water by ultrafiltration, but this does not remove the smallest of organic particles. DBNPA is the most effective biocide, but these aren’t often used and have to be removed from the water. Chlorination is an effective way to remove bacteria, but it can cause degradation of the membrane, and must be removed before the reverse osmosis occurs. In our process we will use UV irradiation to remove any organic material that passes through the microfilters.<ref name = "Desalination"/>

Scaling can reduce the flux across the membrane and can build up in pipe corners, causing backup. It is caused by inorganic salts -- specifically sulfates and calcium and magnesium carbonate -- present in the water. Scaling can be prevented and/or controlled with acids, antiscalants, or ion exchange softening. Antiscalants are good for removing sulfates, but not much else. Acidification will only be effective at removing CaCO3 from water but, the decreased pH may deteriorate the membrane and pipes, so it may require future research.<ref name = "Desalination"/> Ion exchange is very expensive and requires more capital to undergo, but is overall the more effective option.

The ion exchanger is a pretreatment system which acts as a method to fight membrane scaling in reverse osmosis. The ion exchanger works by using sodium carboxylate as a resin which will react with Chlorine and Magnesium ions, removing them from the feed. The sodium is substituted out and enters the feed stream, increasing the overall sodium level and TDS. The lack of magnesium and chlorine ions means that there is nothing to react with the carbonates to form scaling salts on the membranes.

==Desalination (Reverse Osmosis)==

Two pumps (P-105 & P-106) take the adjusted feed seawater up to 58 bar. However, before this pumping occurs, about 70% of the original stream gets separated and heads into a PX Pressure Exchanger. This separated stream later joins together with the pumped stream to enter the reverse osmosis system at the required pressure of 58 bar. The streams can combine together as they have equal pressure. It is the concentrate stream from the output of the reverse osmosis system that transfers its energy to the adjusted feed in the PX Pressure Exchanger.

[[File:Figure_3.jpg|left|300px|frame|Figure 3. Inside look of pressure exchanger]]
The pressure of the concentrate stream is about 57 bar and it is fed into an exchanger which works by using vessel that resembles a small pipe within a rotating wheel. The low pressure (LP) feed enters the vessel at 1 bar (see Figure 3). The feed turns and is pushed out into an outlet pipe by the high pressure (HP) concentrate stream at ~57 bar. The feed is now a HP stream via this exchanged energy. The now LP concentrate stream turns and is pushed out into an outlet pipe by the LP feed entering the vessel.

98% of the energy within the concentrate stream can be recovered via the PX Pressure Exchanger.<ref>Desalination Product and Applications Catalog 2013. “Advance Technologies for Seawater Reverse Osmosis and Brackish Water Applications”. Energy Recovery, Inc. 2012-2014.</ref >Thus, the now HP feed stream is not quite 58 bar and will need to be pumped by a small booster pump (P-107).

When designing a reverse osmosis system, we need to make decisions on what specific kind of reverse osmosis system will be implemented. When making the decision, we investigated the advantages of different types of reverse osmosis among the following options: batch vs. continuous process, number of stages, and number of elements.<ref>FILMTEC Reverse Osmosis Membranes. “Water & Process Solutions Technical Manual.” Dow Chemical Company.</ref> Our decision was made based on process optimization which is described later in this report. The end result of that optimization was to go with a single stage, continuous process with five elements in series.

=Post Treatment=

The permeate is sent to post treatment after the salt content is analyzed and controlled if too high or too low due to any unexpected errors. As stated above, the post treatment process will be done 50 meters from our plant by the Long Beach Water Reclamation Plant. The reason we aren’t performing post treatment “in-house” is due to the increased cost it would take for us to build and run the equipment for post treatment. We have the option to use the government’s water treatment facilities that already exist and our 15 million gallons per day output is a relatively small addition to the large Long Beach Water Reclamation Plant. We will assume any costs to post treatment will be pushed to the consumer in the price of purchasing the drinking water from the Long Beach Water Reclamation Plant.

Post treatment at any facility works to bring water to “taste standards”. The water that is the permeate of our process is already meeting L.A. drinking standards as shown above in Table 1 after our mass balance calculations around the reverse osmosis system (Appendix 3, Table 1). The post treatment increases the pH from 6.3 to 7.3 with the addition of base and then adds minerals like fluoride, calcium, and magnesium back in to improve taste.

==Brine Treatment and Byproduct Waste==

A natural consequence of a reverse osmosis procedure is that a lot of salt will be leftover from the membrane separation in a very concentrated form. This is called “brine” and will have to be taken care of. According to EPA requirements, the salt concentration of water entering the environment (via land or sea) cannot be higher than 40,000 ppm.8 The salt concentration of our brine is much higher (around 70,000 ppm). Thus, the brine must be removed or treated. We decided to treat the brine by diluting it as removing brine can be difficult and may require lots of land (e.g. evaporation ponds).

After the concentrate exchanges energy in the energy exchanger, it becomes a low pressure stream. It is pumped into a dilution tank and mixes with 18.2 million gallons per day of seawater from the inlet to become diluted brine which is pumped 600 meters out to sea.

The waste byproducts are the solid buildup on the filters and membranes. The filters and membranes also have a set lifetime and thus will also need to be removed. The waste is all non-hazardous and thus can be put in the municipal dump.

=Process Details and Equipment Sizing=

==Filters and UV Treatment [Pretreatment]==
The cost of our ultrafiltration system is estimated using data from the University of Tennessee at Knoxville.<ref>Qiang He and R. Bruce Robinson Membrane Filtration Systems. “Membrane Module Costs.” University of Tennessee at Knoxville. December 2014.</ref> It estimated the cost of the cost of the ultrafiltration $0.057 per 1000 gallons and the construction and capital costs are $0.39 per gallon per day. The cost of UV treatment is based off of units available from buyultraviolet.com for units that can treat about 450 gallons per minute.<ref>Atlantic Ultraviolet Corporation. “Megatron Water Disinfection System.”</ref> Due to the water having to be in relatively small pipes for treatment we do not expect the costs to scale very much, so we believe these units to be an accurate cost estimate.

==Ion Exchanger [Pretreatment] ==

The ion exchanger was sized using data from DOW. Our process will use 60 vessels with ion exchange beds of volume 5.73 m3. With a flow rate of 30 bed volumes/hour (10,312.2 m3/hr) the system of beds has a capacity of 125% of the feed. This allows for periodic regeneration of the resin, a weak acid cation resin (DOW MAC-3 Resin). This weak cation resin will remove cationic calcium and magnesium from the water to prevent scaling. The mass flow was calculated via a mass balance on the ion exchanger (Appendix 3).

==Reverse Osmosis System [Desalination]==

Each of the 2607 Reverse Osmosis SW30HRLE-440i [Seawater 30 High Rejection Low Energy 440 ft2 interlocking cap membrane] units have the same dimensions. A, the net length, is 1016 mm. B, the total length including the endcap, is 1029 mm. C, the diameter of the filter element, is 201mm. D, the diameter of the permeate pathway, is 29 mm. For a single stage, we have five of the elements in series so the overall length of a single stage is expected to be roughly 5145 mm. The membrane material is polyamide thin film composite. It will last three years.
Operating Conditions [ROSA 9.0]: 
* Temperature: 30oC 
* Pressure: 58 bar 
* Flow Rate: 52.3 Million gal per day 
* Recovery: 28.7%

[[File:Figure_4.png|center|400px|frame|Figure 4. Diagram of Dow FILMTEC SW30HRLE-440i]]
Cleaning Schedule: 
* Conservative Estimate: Once every three months (three days available for cleaning)<ref name = "Bates"/> 
* DOW FILMTEC Cleaning Conditions 
** The normalized permeate flow drops 10%
* The normalized salt passage increases 5 - 10%

==Seawater Holding Tank and Dilution Tank [Brine Treatment]==

The initial seawater holding tank works as a buffer to prevent surges and help perform shutdowns quickly and easily. We expect to run continuously for 350 days a year with the rest of the available time spread out and used for shutting down the plant for cleaning membranes, filters, and the fixing of unexpected errors. There is a big tradeoff between having a holding tank with high retention time and having a holding tank with low cost. With that in mind, we erred on the side of lowering capital cost and went with a retention time of one hour. The volume could be found by multiplying the retention time with the feed flow rate (Stream 1). We assumed that a good height to width ratio for a large water tank would be 4:1.

A mass balance was done around the dilution tank to find what amount of seawater would be necessary to dilute the brine to 45,000 ppm salinity (Appendix 3; Table 3). With this the feed flow rate into the tank could be found. It was found that it takes about 20 minutes of retention time for good dilution of the brine in seawater.8 We again assumed that a good height to width ratio for a large water tank would be 4:1.

We assumed a corrosion allowance of 4 mm wall thickness. We calculated wall thickness using vessel pressure, hoop stress, longitudinal stress, and weld factor. Each tank is made with SS316 to protect against corrosion.<ref>Towler G, Sinnott R. Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013.</ref>

<center>
'''Table 2.''' Tank sizes.
[[File:Table_2.jpg|center|600px]]
</center>

=Process Optimization=
<center> '''Table 3.'''Optimization Criteria
[[File:Table_3.jpg|center|400px]]
</center>
In order to optimize the reverse osmosis system, numerous simulations were run. Three representative cases are reported as shown above (Table 3). The objective of the optimization was to minimize the required water price to break even in 10 years. In order to achieve this goal, the balance between the capital cost from the number of the filter elements and utility cost from energy consumption was weighed. Before the energy recovery system was implemented, the general trend was that if we have more filter elements in series, overall recovery increases but recovery rate per filter decreases, and we need more filter elements. While increasing the elements in series increase the required number of filters, and thereby increases the capital cost, we require less energy to produce the same amount of water. After the energy recovery system is implemented, the correlation between the number of elements in series and the required energy becomes more complicated because high pressure concentrate is used to pressurize the feed. However, the trade-off between the utility cost and capital cost still exists. Of the three cases shown above, single stage, five elements RO system turned out to be the most optimal solution because single stage, single element system required higher utility cost and two stage system slightly did not meet the LA standard. Detailed information which was used for the optimization is available in Appendix 3 and in Appendix 6.

=Ethical, Environmental, and Societal Considerations=

There are a few key considerations we have followed to meet ethical, environmental, and societal guidelines and requirements: 
* Meet L.A. Drinking Water Standards
** Correct ion concentration
* Moved near water treatment plant
** Pipes will not corrode from low pH
* Working below pressure limits
* Water will be locally served and meet taste standards
** Via Post Treatment mineralization
* Follows EPA Environmental Salinity Requirements
* Reduced Energy Requirements
** PX Pressure Exchanger
** Optimized RO System
** Reverse Osmosis is itself Energy Saving

For further environmental considerations, see effluent summary in Appendix 4.

=Economic Analysis=

The pricing of the majority of the process was completed using the Aspen Economic Analyzer, however we had determine costs for the ultrafiltration, reverse osmosis membranes, and the ultraviolet treatment was completed separately. This led to a total ISBL of $56 MM, OSBL of $22.4 MM, Engineering Costs of $22.4 MM and a Contingency of $7.8 MM. We also chose a working Capital of $10 MM. This working capital is sufficient because we have no raw materials costs. We also had FCOP of $5.2 MM and VCOP of $3.5 MM.

We originally were planning on pricing our water at $700 per acre-foot, but after doing economic analysis we had a 10 year NPV of $-76.2 MM. After determining did a sensitivity analysis and determined that if our water was priced at $2104 per acre-foot we could have an NPV of $0 after 10 years and if we were able to get the same price of $2257 per acre-foot as the plant in Carlsbad we would have a 10 year NPV of $8.3 MM which increases to $41.1 MM at 20 years. For more detailed data about the economic analysis see Appendix 6.

=Conclusion and Recommendations=

Our mission was to create an economically viable process that will desalinate water for human consumption in the L.A. County. We have found after research and optimization that the best process to undergo desalination is a continuous single stage, five elements reverse osmosis system. Our process can meet the required L.A. drinking water standard and we have set up plans for post treatment at a local water reclamation site. While meeting the drinking water standards, we also met our ethical, environmental, and social standards.

Nevertheless, at a sale price of $700 per acre-foot, our proposed desalination plant in Los Angeles, California looks to be a very expensive project to undergo. Without any sort of government assistance to carry the heavy capital costs of such a project, we would suggest the plant not be built. However, after performing a sensitivity analysis at higher (and less conservative) price points we reached a breakeven point in 10 years with a sale price of $2104 per acre-foot of drinking water. Research has shown that the droughts striking California look only to get worse in the near future. With limited freshwater resources, there is no option other than to consider desalination. Therefore it is reasonable to expect we could get near the same price deal as local desalination plants like the Carlsbad Plant ($2257 per acre-foot) As we have met all design constraints for this project, we hope to convince policymakers to adopt our project at least $2104 per acre-foot water. If this deal is reached, we recommend that this project and design be funded.

However, there are ways to improve our design with future work. Right now all the pipes diameters are based on having a flowrate of 1.5 m/s. But in reality, we’d need to use standard piping sizes to get more accurate pricing and energy costs. Also, in reality we would be purchasing some of these vessels in bulk quantities (i.e. bulk filters) so the price could be lowered with negotiation and economies of scale. We also have pumps that are providing very little pressure for their high cost. We can remove these pumps and simply run at higher pressures earlier. We can potentially reduce early energy costs by building a “floating” desalination plant in the sea to reduce pumping distance from the well.

=Appendices=

==Appendix 1: Design Basis==
[[File:Appendix_1.png|thumb|border|center|middle|]]

==Appendix 2: Process Flow Diagram==
[[File:Appendix_2.png|center]]
==Appendix 3: Mass and Energy Balance Tables==
'''Table 1'''. Mass balance on 1 stage 5 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T1.png|center]]

'''Table 2'''. Mass balance on 1 stage 1 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T2.png|center]]

'''Table 3'''. Mass balance on 2 stage 5,5 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T3.png|center]]

'''Table 4'''. 1 stage 5 elements CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T4.png|center]]

'''Table 5'''. 1 stage 1 element CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T5.png|center]]

'''Table 6'''. 2 stages 5,5 elements CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T6.png|center]]

'''Table 7'''. Energy balance on 1 stage 5 elements RO system
[[File:Appendix_3_T7.png|center]]

==Appendix 4: Effluent Summary==
[[File:Appendix_4.png|center]]

==Appendix 5: Cost of Production==
[[File:Appendix_5.png|center]]

==Appendix 6: Economic Anlaysis==
===Case I ===
Water is $700; RO with 1 stage and 5 elements
[[File:Appendix_6_C1.png|center]]

===Case II===
Water is $1400; RO with 1 stage and 5 elements
[[File:Appendix_6_C2.png|center]]

===Case III: Carlsbad ===
Water is $2257; RO with 1 stage and 5 elements
[[File:Appendix_6_C3.png|center]]

===Case IV: Breakeven (NPV = 0 at 10 years)===
Water is $2104; RO with 1 stage and 5 elements
[[File:Appendix_6_C4.png|center]]

===Case V: Breakeven (NPV = 0 at 10 years)===
Water is $2112; RO with 2 stages and 5,5 elements
[[File:Appendix_6_C5.png|center]]

===Case VI: Breakeven (NPV = 0 at 10 years)===
Water is $2138 with 1 stage and 1 element
[[File:Appendix_6_C6.png|center]]

==Appendix 7: ROSA Simulation Outputs==

'''Table 1'''. System detail of 1 stage 5 elements RO system
[[File:Appendix_7_T1.png|center]]

'''Table 2'''. System detail of 1 stage 1 elements RO system
[[File:Appendix_7_T2.png|center]]

'''Table 3'''. System detail of 2 stage 5,5 elements system
[[File:Appendix_7_T3.png|center]]

'''Table 4'''. 1 stage 5 elements RO system concentrations and pH in each stream
[[File:Appendix_7_T4.png|center]]

'''Table 5'''. 1 stage 1 element RO system concentration and pH in each stream
[[File:Appendix_7_T5.png|center]]

'''Table 6'''. 2 stage 5,5 elements RO system concentrations and pH in each stream
[[File:Appendix_7_T6.png|center]]

=References=
<references/>

Desalination - Team E

2016-03-12T03:23:55Z

Bts511: /* Appendix 1: Design Basis */

Group E Corporation

Authors: Hassan Ali, Woo Soo Choe, Brett Sleyster, Jake Stolley

Instructors: Fengqi You, David Wegerer

March 11, 2016

=Executive Summary=

Outlined in this report is the project model report for the proposed desalination plant in Los Angeles, California which will provide drinking water to local citizens using reverse osmosis process. This location was chosen due to the region’s propensity for droughts in recent years. The fresh water resources in California have plummeted as water prices have risen 3-4% each year. This is only expected to continue. Other companies have worked deals with the Californian government to create desalination plants that provide a new fresh water drinking source. This suggests that there is a strong market here with potential growth.

The plant will have an output of 15 million gallons per day, which is about medium size relative to other plants. Thus the build time will be about one year. The plant will run for 30 years, with 350 days per year of use. Shutdown time will be used for cleaning and maintenance of equipment. The plant will be built on Long Beach, 500 meters from the coast nearby the Long Beach Water Reclamation Plant (LBWP). The input will withdraw 70.5 million gallons per day of seawater that will be used for reverse osmosis and dilution of brine.

Reverse osmosis was chosen as the desalination method for its low energy cost and flexibility. We designed a process that will pretreat the seawater to remove waste, colloids, bacteria, and membrane-damaging scalants. Once this seawater is treated is pumped to heavy pressures through a reverse osmosis system. After doing process optimization under various system parameter changes, we found a continuous, single stage, five element reverse osmosis process to be best. It produces a near 30% recovery. We reduce energy costs by including a pressure exchanger between our concentrate outlet and adjusted feed into the reverse osmosis system.

After performing mass and energy balances, we were able to size equipment and find our permeate composition. It turns out our process will meet L.A. standards for drinking water quality. However, post treatment will be done by the local LBWP to make the water taste better and be non-caustic to pipes. This will reduce “in-house” costs. Many environmental considerations were made, to meet our environmental standards. We meet EPA salinity requirements for our brine waste by diluting it. We also don’t use a lot of energy and thus reduce our fossil fuel consumption and CO2 generation. To be safe, we work well under max pressure limits for our equipment. We make sure we are ethical and meet all standards that our local county requires of us.

After an economic evaluation at a conservative well-water price of $700 per acre-foot, we realize that our revenue is way too low to cover our costs. Capital costs are especially large with regards to our large water tanks. We find that we reach a 10 year NPV breakpoint of $2104 per acre-foot after performing a sensitivity analysis. The Carlsbad Plant in San Diego worked out a deal for $2257 per acre-foot.

As we have met all design constraints for this project, we hope to convince policymakers to adopt our project at least $2104 per acre-foot water. If this deal is reached, we recommend that this project and design be funded. There is future work to be done to improve costs such as performing pump optimization. However, this work is minimal and won’t harm our bottom line if we reach our proposed price points.

__TOC__

=Introduction=

Due to a combination of various factors--including rising human population, climate change, increased energy consumption, and land erosion--the availability of fresh water as a resource for human consumption has been dwindling. This has had a negative impact on the world, but this impact has been especially felt in areas with natural shortages in fresh water. Desert regions, such as Saudi Arabia, have resorted to desalinating seawater at the coast and transporting up to their cities. This encompasses their entire freshwater system as very few natural water sources remain. In the United States there is a region which has shown increasing signs of following in the footsteps of Saudi Arabia and becoming fully reliant on desalination: Los Angeles, California.

=Location and Market Analysis=

In this project, Los Angeles was chosen as the location our plant, because the area has been subject to a severe drought for the last four years, as shown in Figure 1. The rapid decrease in the available water has caused permanent damages to the ground, decreased the amount of water in the reservoir from 45% to 25%, and reduced the amount of hydroelectric power produced in the state. Such deficiency in usable water has increased the water prices in Los Angeles by 6-7% annually for the last decade and is expected to rise by 3.4% annually for typical customers for the next five years. For customers who use more than 20000 gallons a year, the bill is expected to increase by 34% by 2021. This price increase is occurring because of both drought and maintenance. <ref>Pacific Institute: The California Drought. 2014.</ref>

[[File:Figure_1.jpg|right|200px|frame|Figure 1. California Drought Index.]]
As of March 2015 the Metropolitan Water District of Southern California was willing to pay as much as $700 per acre-foot of water to farmers in the Sacramento Valley, so we assume this will be the target price of the water produced by the this desalination plant.<ref>Hessel, Phil. “Dry Southern California Offers Northern Farmers Top Dollar for Water”. NBC News. March 18, 2015. </ref>

Currently, there is a plant in Carlsbad which supplies drinking water to the San Diego area. The plant can produce up to 50 million gallons per day and is the nation's largest water desalination plant. The water from this plant costs $2257 per acre-foot for the first 48000 acre-feet and about $2000 for any additional acre-foot. The county signed a contract with this plant to purchase the first 48000 acre-feet of water for the next 30 years.<ref>Carlsbad Desalination Plant. “Project Overview”. January 2016.</ref> Also, there is another desalination plant in Tampa that produces about 25 million gallons of water per day, and over 2000 other public and industrial plants with production capacity of over 300000 gallons per day. When conducting the economic analysis our plant, such numbers presented by other plants must be considered, so our project can yield positive profit.<ref>Leven, Rachel. “U.S. Desalination Industry Grows Since 2000; See as Essential to Meeting Supply Needs”. Bloomberg BNA. August 21, 2013.</ref>

The profitability of the project is naturally heavily dependent on the price of water in California. If the drought ends, water supply will increase and the water price will drop significantly, but if the drought continues, the price of water in California could rise even more rapidly and allow desalination plants to be more profitable.<ref>Southern California Public Radio. “Where is California water use decreasing?”. December 2015.</ref> It may be difficult to make this plant profitable without a guaranteed purchase or other subsidy from the local governments because the presence of pre-existing plants maintains the water price relatively low. However, external sources of profit such as guaranteed purchase or subsidy are expected because the ramifications of not having desalination plant or other sources of water may be significantly more expensive than the subsidizing a desalination plant assuming the drought continues.<ref>Weiser, Matt. “Could desalination solve California’s water problem?”. The Sacramento Bee. October 18, 2014.</ref>

=Design Basis and Technical Approach=

[[File:Figure_2.jpg|left|200px|frame|Figure 2. Vertical Well Feed]]
Considering the number of pre-existing desalination plant and the potential risk of the drought ending, our plant will be tentatively designed to produce 15 million gallons of drinkable water per day to homes and businesses in the Los Angeles county area. With a recovery rate of near 30%, and a requirement for diluting any brine, the input feed flowrate will be 70.5 million gallons per day. Following the similarly built Carlsbad plant, we expect the plant to last 30 years. The project building timeline is one year based on our relatively small plant size. We will run for 350 days a year while shutting down for cleaning and maintenance once every 3 months (based on DOW standards).<ref name = "Bates">Bates, Wayne. “Cleaning your RO”. Hydranautics. </ref> The plant will be built 500 meters from the coastline on Long Beach. A borehole will be drilled 10 meters into the ground. At this depth there should be seawater aquifers below the coast that regularly recharges itself with seawater from the ocean. A vertical well containing intake holes will be put into the borehole. The intake holes will be equipped with large size filters that can easily be replaced. At the surface, the well will be connected to a large-scale pump in a well pump house. From there, the seawater will be horizontally pumped into a seawater tank in the desalination plant. This will be considered the start of our process. The process ends after pretreatment, desalination, post treatment, and brine treatment are completed. Appendix 1 shows more details on the Design Basis.

The feed composition is assumed to be the standard seawater “average” based on the DOW Chemical ROSA 9.0 Technical Manual in Table 1. However, in reality the seawater composition at the plant site will have to be chemically tested. The final product will be based on L.A. standards for drinking water quality (Table 1). The specifications in the table are upper limits that we cannot exceed. We will aim to reach the salt level standards but any remineralization and pH treatment will be exported and paid for due to the added complexity and cost in making such a system in-house.

Due to the acidity of the desalinated water coming out of the reverse osmosis (RO) system (pH ~ 6), the pumps and pipes carrying the desalinated water for post treatment will likely corrode over time. We have decided to place the desalination plant by the Long Beach Water Reclamation Plant (about 50 meters away) to prevent pipe corrosion.

[[File:Table_1.jpg|center|200px|frame|Table 1. Concentration of the standard seawater, the output from the RO system, and LA standard.]]

=Desalination Overview and Alternatives=

We first began our research into the current landscape of the desalination industry, particularly of plants that were built in the United States. We discovered there were many processes but the main ones included: 
* Membrane Separation (Forward Osmosis and Reverse Osmosis)
* Thermal Separation (Multi Stage Flash Distillation)
* Phase Change Separation (Freeze Separation)
Each process was analyzed for their pros and cons. For our project design, we decided to go with reverse osmosis as our desalination process. This is because it is the most commonly used process in the industry, and for good reason. Unlike thermal separation, which requires huge amounts of energy to supply heat, reverse osmosis is very energy efficient. This reduces energy costs that would be spread across the plant’s lifetime.<ref name = "Republic">Republic of Mauritius NY. "Technical Aspects of Desalination Plant". Sustainable Sanitation and Water Management. </ref>

Forward osmosis is a newer technology being developed that also uses membranes to separate water and salt and may be more energy efficient overall than reverse osmosis. This is because forward osmosis doesn’t require a pump to pressurize seawater through the membranes, but rather leaves a higher salt concentrated solution (AKA draw solution) on the other side of the membrane for the seawater to be drawn into. Then a second step is required to regenerate the concentrated draw solution and produce purified water.<ref>Mallinson, Alissa. "Study Shows Forward Osmosis Desalination Not Energy Efficient." MIT News [Boston]. July 23, 2014.</ref> We decided to move away from forward osmosis because it is a process that is still being refined and developed for desalination and lacks the data, programs, and technologies that reverse osmosis users benefit from. Additionally, a recent study from MIT has claimed that due to the inherent nature of forward osmosis requiring that seawater be drawn to a higher concentrated salt solution before the water and salt is separated, it cannot possibly be more energy efficient than reverse osmosis.<ref>O'Hern, Sean C., Karnik, Rohit. Nanofiltration across Defect-Sealed Nanoporous Monolayer Graphene. American Chemical Society 15.5 (2014); 3254-260. </ref>

Other processes such as freeze separation were too restrictive in their feed and process demands and required a lot of capital. Ultimately, reverse osmosis is simply the best process that exists for desalination and has been ranked the best when performance is evaluated in terms of technical, environmental, and economic aspects. It is a process that is flexible in water quantity, operation start-up and shut-off, and site location.

Nevertheless there are disadvantages to reverse osmosis that must be considered. The feed seawater must not be too salty and too low quality. The pressure to pump the seawater is enormous. The construction time is long for large plants. The capital costs are also high when compared with other processes such as thermal separation. But all of these disadvantages are ultimately outweighed by the advantages reverse osmosis provides.<ref name = "Republic"/>

=Process Overview and Alternatives=

See Appendix 2 (Figure 1) for the Process Flow Diagram (PFD). The process itself is detailed below. The stream flow rates are in Appendix 2 (Table 1). We assumed 10 meters of pipe for every stream except the inlet and outlet streams of the entire PFD.

==Pretreatment==

After the 10 meter vertical well pumps seawater from the Long Beach coast, the feed is pumped 500 meters to a seawater tank that makes it easy to shut-down the system and reduce the large volume of flow. The level control keeps the seawater from overfilling the tank. From there, part of the seawater goes to brine treatment while the rest go to pretreatment. The purpose of pretreatment in the desalination process is to prevent equipment damage by removing particles and microorganisms from the water. A variety of pollutants -- including calcium carbonate, sulfates, silica, clay, bacteria, and other colloids and microorganisms -- can disrupt the reverse osmosis process. Media filtration uses fine grains such as sand or anthracite to filter the particles. We used a large filter to catch larger particle material (e.g. sand). However, simple filtration is not enough to prevent colloid fouling. An ultrafilter (i.e. microfiltration) is used to remove virtually all suspended matter, effectively preventing colloid fouling. The costs associated with ultrafiltration are not prohibitive in terms of the project’s budget.<ref name = "Desalination">Desalination Pre-treatment. Lenntech.</ref>

Reverse osmosis systems can also be disrupted by the growth of microbiological films of bacteria on the film and pipe walls. They are typically treated as colloids during pretreatment, and most organisms are removed from the water by ultrafiltration, but this does not remove the smallest of organic particles. DBNPA is the most effective biocide, but these aren’t often used and have to be removed from the water. Chlorination is an effective way to remove bacteria, but it can cause degradation of the membrane, and must be removed before the reverse osmosis occurs. In our process we will use UV irradiation to remove any organic material that passes through the microfilters.<ref name = "Desalination"/>

Scaling can reduce the flux across the membrane and can build up in pipe corners, causing backup. It is caused by inorganic salts -- specifically sulfates and calcium and magnesium carbonate -- present in the water. Scaling can be prevented and/or controlled with acids, antiscalants, or ion exchange softening. Antiscalants are good for removing sulfates, but not much else. Acidification will only be effective at removing CaCO3 from water but, the decreased pH may deteriorate the membrane and pipes, so it may require future research.<ref name = "Desalination"/> Ion exchange is very expensive and requires more capital to undergo, but is overall the more effective option.

The ion exchanger is a pretreatment system which acts as a method to fight membrane scaling in reverse osmosis. The ion exchanger works by using sodium carboxylate as a resin which will react with Chlorine and Magnesium ions, removing them from the feed. The sodium is substituted out and enters the feed stream, increasing the overall sodium level and TDS. The lack of magnesium and chlorine ions means that there is nothing to react with the carbonates to form scaling salts on the membranes.

==Desalination (Reverse Osmosis)==

Two pumps (P-105 & P-106) take the adjusted feed seawater up to 58 bar. However, before this pumping occurs, about 70% of the original stream gets separated and heads into a PX Pressure Exchanger. This separated stream later joins together with the pumped stream to enter the reverse osmosis system at the required pressure of 58 bar. The streams can combine together as they have equal pressure. It is the concentrate stream from the output of the reverse osmosis system that transfers its energy to the adjusted feed in the PX Pressure Exchanger.

[[File:Figure_3.jpg|left|300px|frame|Figure 3. Inside look of pressure exchanger]]
The pressure of the concentrate stream is about 57 bar and it is fed into an exchanger which works by using vessel that resembles a small pipe within a rotating wheel. The low pressure (LP) feed enters the vessel at 1 bar (see Figure 3). The feed turns and is pushed out into an outlet pipe by the high pressure (HP) concentrate stream at ~57 bar. The feed is now a HP stream via this exchanged energy. The now LP concentrate stream turns and is pushed out into an outlet pipe by the LP feed entering the vessel.

98% of the energy within the concentrate stream can be recovered via the PX Pressure Exchanger.<ref>Desalination Product and Applications Catalog 2013. “Advance Technologies for Seawater Reverse Osmosis and Brackish Water Applications”. Energy Recovery, Inc. 2012-2014.</ref >Thus, the now HP feed stream is not quite 58 bar and will need to be pumped by a small booster pump (P-107).

When designing a reverse osmosis system, we need to make decisions on what specific kind of reverse osmosis system will be implemented. When making the decision, we investigated the advantages of different types of reverse osmosis among the following options: batch vs. continuous process, number of stages, and number of elements.<ref>FILMTEC Reverse Osmosis Membranes. “Water & Process Solutions Technical Manual.” Dow Chemical Company.</ref> Our decision was made based on process optimization which is described later in this report. The end result of that optimization was to go with a single stage, continuous process with five elements in series.

=Post Treatment=

The permeate is sent to post treatment after the salt content is analyzed and controlled if too high or too low due to any unexpected errors. As stated above, the post treatment process will be done 50 meters from our plant by the Long Beach Water Reclamation Plant. The reason we aren’t performing post treatment “in-house” is due to the increased cost it would take for us to build and run the equipment for post treatment. We have the option to use the government’s water treatment facilities that already exist and our 15 million gallons per day output is a relatively small addition to the large Long Beach Water Reclamation Plant. We will assume any costs to post treatment will be pushed to the consumer in the price of purchasing the drinking water from the Long Beach Water Reclamation Plant.

Post treatment at any facility works to bring water to “taste standards”. The water that is the permeate of our process is already meeting L.A. drinking standards as shown above in Table 1 after our mass balance calculations around the reverse osmosis system (Appendix 3, Table 1). The post treatment increases the pH from 6.3 to 7.3 with the addition of base and then adds minerals like fluoride, calcium, and magnesium back in to improve taste.

==Brine Treatment and Byproduct Waste==

A natural consequence of a reverse osmosis procedure is that a lot of salt will be leftover from the membrane separation in a very concentrated form. This is called “brine” and will have to be taken care of. According to EPA requirements, the salt concentration of water entering the environment (via land or sea) cannot be higher than 40,000 ppm.8 The salt concentration of our brine is much higher (around 70,000 ppm). Thus, the brine must be removed or treated. We decided to treat the brine by diluting it as removing brine can be difficult and may require lots of land (e.g. evaporation ponds).

After the concentrate exchanges energy in the energy exchanger, it becomes a low pressure stream. It is pumped into a dilution tank and mixes with 18.2 million gallons per day of seawater from the inlet to become diluted brine which is pumped 600 meters out to sea.

The waste byproducts are the solid buildup on the filters and membranes. The filters and membranes also have a set lifetime and thus will also need to be removed. The waste is all non-hazardous and thus can be put in the municipal dump.

=Process Details and Equipment Sizing=

==Filters and UV Treatment [Pretreatment]==
The cost of our ultrafiltration system is estimated using data from the University of Tennessee at Knoxville.<ref>Qiang He and R. Bruce Robinson Membrane Filtration Systems. “Membrane Module Costs.” University of Tennessee at Knoxville. December 2014.</ref> It estimated the cost of the cost of the ultrafiltration $0.057 per 1000 gallons and the construction and capital costs are $0.39 per gallon per day. The cost of UV treatment is based off of units available from buyultraviolet.com for units that can treat about 450 gallons per minute.<ref>Atlantic Ultraviolet Corporation. “Megatron Water Disinfection System.”</ref> Due to the water having to be in relatively small pipes for treatment we do not expect the costs to scale very much, so we believe these units to be an accurate cost estimate.

==Ion Exchanger [Pretreatment] ==

The ion exchanger was sized using data from DOW. Our process will use 60 vessels with ion exchange beds of volume 5.73 m3. With a flow rate of 30 bed volumes/hour (10,312.2 m3/hr) the system of beds has a capacity of 125% of the feed. This allows for periodic regeneration of the resin, a weak acid cation resin (DOW MAC-3 Resin). This weak cation resin will remove cationic calcium and magnesium from the water to prevent scaling. The mass flow was calculated via a mass balance on the ion exchanger (Appendix 3).

==Reverse Osmosis System [Desalination]==

Each of the 2607 Reverse Osmosis SW30HRLE-440i [Seawater 30 High Rejection Low Energy 440 ft2 interlocking cap membrane] units have the same dimensions. A, the net length, is 1016 mm. B, the total length including the endcap, is 1029 mm. C, the diameter of the filter element, is 201mm. D, the diameter of the permeate pathway, is 29 mm. For a single stage, we have five of the elements in series so the overall length of a single stage is expected to be roughly 5145 mm. The membrane material is polyamide thin film composite. It will last three years.
Operating Conditions [ROSA 9.0]: 
* Temperature: 30oC 
* Pressure: 58 bar 
* Flow Rate: 52.3 Million gal per day 
* Recovery: 28.7%

[[File:Figure_4.png|center|400px|frame|Figure 4. Diagram of Dow FILMTEC SW30HRLE-440i]]
Cleaning Schedule: 
* Conservative Estimate: Once every three months (three days available for cleaning)<ref name = "Bates"/> 
* DOW FILMTEC Cleaning Conditions 
** The normalized permeate flow drops 10%
* The normalized salt passage increases 5 - 10%

==Seawater Holding Tank and Dilution Tank [Brine Treatment]==

The initial seawater holding tank works as a buffer to prevent surges and help perform shutdowns quickly and easily. We expect to run continuously for 350 days a year with the rest of the available time spread out and used for shutting down the plant for cleaning membranes, filters, and the fixing of unexpected errors. There is a big tradeoff between having a holding tank with high retention time and having a holding tank with low cost. With that in mind, we erred on the side of lowering capital cost and went with a retention time of one hour. The volume could be found by multiplying the retention time with the feed flow rate (Stream 1). We assumed that a good height to width ratio for a large water tank would be 4:1.

A mass balance was done around the dilution tank to find what amount of seawater would be necessary to dilute the brine to 45,000 ppm salinity (Appendix 3; Table 3). With this the feed flow rate into the tank could be found. It was found that it takes about 20 minutes of retention time for good dilution of the brine in seawater.8 We again assumed that a good height to width ratio for a large water tank would be 4:1.

We assumed a corrosion allowance of 4 mm wall thickness. We calculated wall thickness using vessel pressure, hoop stress, longitudinal stress, and weld factor. Each tank is made with SS316 to protect against corrosion.<ref>Towler G, Sinnott R. Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013.</ref>

<center>
'''Table 2.''' Tank sizes.
[[File:Table_2.jpg|center|600px]]
</center>

=Process Optimization=
<center> '''Table 3.'''Optimization Criteria
[[File:Table_3.jpg|center|400px]]
</center>
In order to optimize the reverse osmosis system, numerous simulations were run. Three representative cases are reported as shown above (Table 3). The objective of the optimization was to minimize the required water price to break even in 10 years. In order to achieve this goal, the balance between the capital cost from the number of the filter elements and utility cost from energy consumption was weighed. Before the energy recovery system was implemented, the general trend was that if we have more filter elements in series, overall recovery increases but recovery rate per filter decreases, and we need more filter elements. While increasing the elements in series increase the required number of filters, and thereby increases the capital cost, we require less energy to produce the same amount of water. After the energy recovery system is implemented, the correlation between the number of elements in series and the required energy becomes more complicated because high pressure concentrate is used to pressurize the feed. However, the trade-off between the utility cost and capital cost still exists. Of the three cases shown above, single stage, five elements RO system turned out to be the most optimal solution because single stage, single element system required higher utility cost and two stage system slightly did not meet the LA standard. Detailed information which was used for the optimization is available in Appendix 3 and in Appendix 6.

=Ethical, Environmental, and Societal Considerations=

There are a few key considerations we have followed to meet ethical, environmental, and societal guidelines and requirements: 
* Meet L.A. Drinking Water Standards
** Correct ion concentration
* Moved near water treatment plant
** Pipes will not corrode from low pH
* Working below pressure limits
* Water will be locally served and meet taste standards
** Via Post Treatment mineralization
* Follows EPA Environmental Salinity Requirements
* Reduced Energy Requirements
** PX Pressure Exchanger
** Optimized RO System
** Reverse Osmosis is itself Energy Saving

For further environmental considerations, see effluent summary in Appendix 4.

=Economic Analysis=

The pricing of the majority of the process was completed using the Aspen Economic Analyzer, however we had determine costs for the ultrafiltration, reverse osmosis membranes, and the ultraviolet treatment was completed separately. This led to a total ISBL of $56 MM, OSBL of $22.4 MM, Engineering Costs of $22.4 MM and a Contingency of $7.8 MM. We also chose a working Capital of $10 MM. This working capital is sufficient because we have no raw materials costs. We also had FCOP of $5.2 MM and VCOP of $3.5 MM.

We originally were planning on pricing our water at $700 per acre-foot, but after doing economic analysis we had a 10 year NPV of $-76.2 MM. After determining did a sensitivity analysis and determined that if our water was priced at $2104 per acre-foot we could have an NPV of $0 after 10 years and if we were able to get the same price of $2257 per acre-foot as the plant in Carlsbad we would have a 10 year NPV of $8.3 MM which increases to $41.1 MM at 20 years. For more detailed data about the economic analysis see Appendix 6.

=Conclusion and Recommendations=

Our mission was to create an economically viable process that will desalinate water for human consumption in the L.A. County. We have found after research and optimization that the best process to undergo desalination is a continuous single stage, five elements reverse osmosis system. Our process can meet the required L.A. drinking water standard and we have set up plans for post treatment at a local water reclamation site. While meeting the drinking water standards, we also met our ethical, environmental, and social standards.

Nevertheless, at a sale price of $700 per acre-foot, our proposed desalination plant in Los Angeles, California looks to be a very expensive project to undergo. Without any sort of government assistance to carry the heavy capital costs of such a project, we would suggest the plant not be built. However, after performing a sensitivity analysis at higher (and less conservative) price points we reached a breakeven point in 10 years with a sale price of $2104 per acre-foot of drinking water. Research has shown that the droughts striking California look only to get worse in the near future. With limited freshwater resources, there is no option other than to consider desalination. Therefore it is reasonable to expect we could get near the same price deal as local desalination plants like the Carlsbad Plant ($2257 per acre-foot) As we have met all design constraints for this project, we hope to convince policymakers to adopt our project at least $2104 per acre-foot water. If this deal is reached, we recommend that this project and design be funded.

However, there are ways to improve our design with future work. Right now all the pipes diameters are based on having a flowrate of 1.5 m/s. But in reality, we’d need to use standard piping sizes to get more accurate pricing and energy costs. Also, in reality we would be purchasing some of these vessels in bulk quantities (i.e. bulk filters) so the price could be lowered with negotiation and economies of scale. We also have pumps that are providing very little pressure for their high cost. We can remove these pumps and simply run at higher pressures earlier. We can potentially reduce early energy costs by building a “floating” desalination plant in the sea to reduce pumping distance from the well.

=Appendices=

==Appendix 1: Design Basis==
[[File:Appendix_1.png|thumb|border|center|middle|]

==Appendix 2: Process Flow Diagram==
[[File:Appendix_2.png|center]]
==Appendix 3: Mass and Energy Balance Tables==
'''Table 1'''. Mass balance on 1 stage 5 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T1.png|center]]

'''Table 2'''. Mass balance on 1 stage 1 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T2.png|center]]

'''Table 3'''. Mass balance on 2 stage 5,5 element RO system, where "out" is the sum of all the concentrate and permeate.
[[File:Appendix_3_T3.png|center]]

'''Table 4'''. 1 stage 5 elements CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T4.png|center]]

'''Table 5'''. 1 stage 1 element CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T5.png|center]]

'''Table 6'''. 2 stages 5,5 elements CO3, HCO3, and CO2 ion mass balance
[[File:Appendix_3_T6.png|center]]

'''Table 7'''. Energy balance on 1 stage 5 elements RO system
[[File:Appendix_3_T7.png|center]]

==Appendix 4: Effluent Summary==
[[File:Appendix_4.png|center]]

==Appendix 5: Cost of Production==
[[File:Appendix_5.png|center]]

==Appendix 6: Economic Anlaysis==
===Case I ===
Water is $700; RO with 1 stage and 5 elements
[[File:Appendix_6_C1.png|center]]

===Case II===
Water is $1400; RO with 1 stage and 5 elements
[[File:Appendix_6_C2.png|center]]

===Case III: Carlsbad ===
Water is $2257; RO with 1 stage and 5 elements
[[File:Appendix_6_C3.png|center]]

===Case IV: Breakeven (NPV = 0 at 10 years)===
Water is $2104; RO with 1 stage and 5 elements
[[File:Appendix_6_C4.png|center]]

===Case V: Breakeven (NPV = 0 at 10 years)===
Water is $2112; RO with 2 stages and 5,5 elements
[[File:Appendix_6_C5.png|center]]

===Case VI: Breakeven (NPV = 0 at 10 years)===
Water is $2138 with 1 stage and 1 element
[[File:Appendix_6_C6.png|center]]

==Appendix 7: ROSA Simulation Outputs==

'''Table 1'''. System detail of 1 stage 5 elements RO system
[[File:Appendix_7_T1.png|center]]

'''Table 2'''. System detail of 1 stage 1 elements RO system
[[File:Appendix_7_T2.png|center]]

'''Table 3'''. System detail of 2 stage 5,5 elements system
[[File:Appendix_7_T3.png|center]]

'''Table 4'''. 1 stage 5 elements RO system concentrations and pH in each stream
[[File:Appendix_7_T4.png|center]]

'''Table 5'''. 1 stage 1 element RO system concentration and pH in each stream
[[File:Appendix_7_T5.png|center]]

'''Table 6'''. 2 stage 5,5 elements RO system concentrations and pH in each stream
[[File:Appendix_7_T6.png|center]]

=References=
<references/>

Separation processes

2016-02-22T05:20:15Z

Bts511: /* Cyclones */

Authors: Nick Pinkerton, [2014] Karen Schmidt, [2014] James Xamplas, [2014] Emm Fulk, [2015] and Erik Zuehlke, [2015] John Dombrowski [2016] , Brett Sleyster [2016] , Robert Cignoni [2016] , and Osman Jamil [2016] 

Stewards: David Chen, Jian Gong, and Fengqi You

Date Presented: February 9, 2014 /Date Revised: February 1, 2014

 

==Introduction==
Essentially all chemical processes require the presence of a separation stage. Most chemical plants comprise of a reactor surrounded by many separators. Separators have a countless number of jobs inside of a chemical plant. A separator can process raw materials prior to the reaction, remove incondensable gases, remove undesired side products, purify a product stream, recycle materials back into the process, and many other jobs that are essential to the process.

Chemical engineers must understand the science of separation and the variety of ways that separation can take place. There are many ways to perform a separation some of these including: distillation, absorption, stripping, and extraction. The science of separation revolves around the presence of two phases that are in contact and equilibrium (Wankat, 2012).

[[File:Sepmeth.JPG|frame|Figure 1. Separation methods by property]]

==Theory==
===Vapor-Liquid Equilibrium===
Separation processes are based on the theory of vapor-liquid equilibrium. This theory states that streams leaving a stage in a separation process are in equilibrium with one another. The idea of equilibrium revolves around the idea that when there is vapor and liquid in contact with one another they are in constantly vaporizing and condensing. Different components in the mixture will condense and vaporize at different rates. There are three types of equilibrium conditions that can be subdivided into thermal, mechanical and chemical potential categories. These separate equilibrium states are given as:

<math>T_{liquid} = T_{vapor}</math>

<math>p_{liquid} = p_{vapor}</math>

<math>chemical potential_{liquid} = chemical potential_{vapor}</math>

==Distillation==
===Flash Distillation===
Flash Distillation is one of the simpler separation processes to be employed in a chemical plant. The main premise of flash distillation is that a portion of a liquid feed stream vaporizes in a flash chamber or a vapor feed condenses. Vapor-liquid equilibrium will cause the vapor phase and the liquid phase to have different compositions. The more volatile component of the mixture will compose of a larger portion of the vapor. This simple separation is easy to manufacture but does not result in large degrees of separation.

Flash distillation requires a feed stream that is pressurized and heated and then passed through a valve into a flash drum. The large pressure drop across the valve will result in a partial vaporization of the fluid. Vapor will be removed overhead from the flash drum while the remaining liquid will collect at the bottom of the drum and be removed. Most flash drums will contain an entrainment eliminator which is a screen that prevents liquid from being carried into the vapor effluent. Figure 2 shows a simple overview of the flash distillation process. As shown, there is a heater that flows into a let-down valve where the two-phase flow begins. Variables y and x are the mole fractions of the more volatile component in the vapor and liquid effluents, respectively.

[[File:Flash.gif|center|frame|Figure 2. Flash Distillation Flow Diagram]]

===Column Distillation===
Distillation columns are the most widely used separation technique used in the chemical industry, accounting for approximately 90% of all separations (Wankat, 2012). Distillations in columns consist of multiple trays that each act at their own equilibrium conditions. Large columns are able to perform complete separations of binary mixtures as well as more complex multi-component mixtures.

[[File:column.jpg|250px|center|]]
===Stages===
Columns are separated into stages by the presence of trays. These trays allow for vapor-liquid contact and equilibrium to occur. Typically, the more stages in a column, the larger separation that can be achieved. There are many different types of trays that can be used in a column.
====Sieve Trays====
The simplest and least expensive tray type is the sieve tray which is a sheet of metal with holes punched into it to allow vapor flow. Sieve trays can have different hole patterns and sizes that will affect the tray efficiency and flow rates.

[[File:sieve.jpg|200px|center|]]

====Sieve Tray Design Procedure====

The design of these plates is done through a trial-and-error process. Most commercial process simulations (such as HYSYS) have default tray designs, and automatically specify dimensions. However, these dimensions selected or calculated by the simulations may not give the best performance for your system, so it is valuable to understand how to design the sieve trays and how specific parameters may affect performance. Hand calculations using the following methods can be used to guide the simulation programs to better design. This section will use sample data to work through an example of the process. The following is a general list of steps for designing a sieve plate:

=====1. Calculate the maximum and minimum vapor and liquid flow rates for the turndown ratio required.=====
This data can be collected from a McCabe-Thiele diagram and/or from process simulation data.

Data from McCabe Thiele diagram, for example: 
Number of stages = 10 
Slope of top operating line = 0.185 
Slope of bottom operating line = 1.43 
Top composition = 98.8 mol% acetone 
Bottom composition = 4 mol% acetone (a) 
Minimum reflux ratio = 0.31 

Collect data rom mass balances: 
MW of feed = (0.6 kmol a/kmol)*(58 kg/kmol a) + (0.4 kmol acetic acid/kmol)*(60 kg/kmol aa) = 59 kg/kmol 
Feed rate = 100 kmol/h 
Using material balances... D = 57.3 kmol/h ; B = 42.7 kmol/h 
Vapor rate, V = D(1+R) = 57.3 kmol/h (1+.31) = 75.1 kmol/h 
Below feed line, slope of BOL = 1.43 = L'm/V'm (liquid rate/vapor rate) 
-- Also, V'm = L'm - B (mass balance on bottom of column) 
-- V'm= 99.3 kmol/h and L'm = 142.0 kmol/h

=====2. Collect or estimate the system physical properties.=====
Here it is important to know information about both the top and bottom of the column. Useful information includes temperature, pressure, column pressure drop (a common assumption is 100 mmH2O per plate), densities, molecular weights, surface tensions, and number of stages (which can be estimated from the McCabe-Thiele diagram).

=====3. Select a Trial Plate Spacing=====
The plate spacing will depend on the column diameter and operating conditions. Plate spacings from 0.15 m to 1.0 m are typically used. The smaller the diameter, the smaller the spacing. Small columns will use close spacing. Columns with diameters above 1.0 m, plate spacings of 0.3 m to 0.6 m are normally used. A good initial estimate is 0.5 m.
[[File:trayspacing.jpg|400px|center|]]

=====4. Estimate the column diameter, based on flooding considerations.=====
Vapor and liquid flow rates will vary along the column, so plate design needs to be considered both above and below the feed. Using plate spacing and FLV (which is the square root of the ratio of the liquid to vapor flow rates), you can obtain the value of K from the plot.

[[File:floodingplot.jpg|400px|center|frame|Figure. Plate Spacing]]
There is a range of vapor and liquid flow rates in which the column needs to be operated. Too low or too high of rates can result in various inefficiencies in the column operation, as shown in the figure below. For example, if the vapor rate is too high, flooding will occur. However, it is not safe to operate on the flooding line. Instead, columns are typically designed for 80% of flooding at the maximum flow rate.
[[File:vap_rate_vs_liq_rate.jpg|400px|center|frame|Figure. Tray behavior]]

=====5. Decide the liquid flow arrangement.=====
Common flow arrangements are single pass (cross flow), double pass, and reverse flow. Using conditions at the bottom of the column, calculate the max volumetric flow rate. Use this flow rate and the column diameter to determine the preferred flow arrangement from the chart below.
[[File:Liquidflow.jpg|200px|center|]]

=====6. Make a trial plate layout: downcomer area, active area, hole area, hole size, weir height.=====
Standard sizes for trays -- and good assumptions for the first iteration -- are: weir height, hw = 50mm ; hole diameter, Dh = 5mm ; plate thickness, tpl = 5mm. From the graph below, the ratio of downcomer area (Ad) to column cross-sectional area (Ac) can be determined from the ratio of weir length (lw) to column diameter (Dc) and vice versa.
[[File:platelayout.jpg|200px|center|frame|Figure. Plate Dimensions]]

=====7. Check the weeping rate=====
Compare the actual vapor velocity to the minimum vapor velocity -- if velocity is too low fluid will "weep" through the tray holes. If the weeping rate is unsatisfactory, return to step 6 and choose different values for the plate layout dimensions. From the chart in step 4, it can be seen that there is a minimum vapor flow rate below which the liquid "weeps" from the tray above.

For the remaining steps in this design process, it is recommended to check your assumptions after each step and revise them as necessary in order to maintain operation in the "sweet spot" of the vapor rate vs. liquid rate plot. Additional iterations may be required as you move through the procedure.

Calculate the maximum liquid flow rate. Calculate the minimum liquid flow rate at 70% turndown (recommended). Calculate the height over the weir as
<math>h_o=750[\frac{L_w}{p_Ll_w}]^\frac{2}{3}</math>

=====8. Check the plate pressure drop=====
<dfn>If the pressure drop calculated here is too high, return to step 6.</dfn>

Proceed to step 9 if the pressure drop assumption is valid.
=====9. Check the downcomer backup. =====
<dfn>If the downcomer backup is too high, return to step 6 or 3.</dfn>

The plate spacing affects the amount of fluid in the downcomer. Calculate the level in the downcomer and the residence time of the fluid to see if the values are valid. Note that residence times greater than 3 seconds are acceptable.

Proceed to step 10 if residence time is acceptable.
=====10. Decide plate layout details.=====
Determine calming zones, the unperforated areas at the inlet and outlet sides of the plate. The width of each zone is usually made the same. Recommended values are: below 1.5 m diameter, 75 mm; above, 100 mm. The unperforated area can be calculated from plate geometry. Also check the hole pitch, or the distance between hole centers. It should not be less than 2.0 hole diameters. A normal range is between 2.5 and 4.0 hole diameters. The shape must also be specified. Square and equilateral triangle holes are used.

=====11. Recalculate the percentage flooding based on the chosen column diameter.=====
An assumption of 80% flooding was chosen so that operation would occur in the "sweet spot." This assumption must be checked by calculating the flooding percentage for a given column diameter.
uv = (max volumetric flow rate)/(net area)

<math>%flooding = \frac{u_v}{u_f}</math>

If the hole pitch is unsatisfactory, return to step 6.
=====12. Check entrainment=====
''If too high, return to step 4'' Use the graph below to determine entrainment from FLV.
[[File:entrainment.jpg|400px|center|]]

The value for fractional entrainment can be used to re-estimate the column efficiency, and reevaluate the number of trays needed. Can return to step 1 for more accurate estimates.

=====13. Optimize design.=====
After returning to step 1 to reevaluate the number of trays, it is valuable to repeat steps 2 through 12 to find the smallest diameter and plate spacing acceptable at the lowest cost.

=====14. Finalize the design.=====
Optional: draw up the plate specification and sketch the layout of the plate.

====Bubble-Cap Trays====
Bubble-cap trays consist of a weir around each hole in the tray which is covered with a cap that has holes or slots to allow vapor passage. Entrainment is about three times larger than a sieve tray. Bubble-cap trays require larger tray spacing than sieve tray design. Bubble-cap trays have been known to have problems with coking, polymer formation, or high fouling mixtures. Recently, very few new bubble-cap columns are being built due to the expense and marginal benefits. However, engineers will likely encounter bubble-cap columns still currently in operation.

====Flow Patterns====
Cross flow columns are the most common pattern for distillation columns. For liquid flows between 50 and 500 Gal/min, a cross flow column is appropriate. When liquid flow is increased above 500 Gal/min, an engineer should consider designing a double pass or multi-pass column. This will reduce the liquid gradient on the tray and reduce the downcomer loading (Wankat, 2012).

===Column Sizing===
Column height will be dependent on the amount of trays required and the spacing between the trays. Normally, tray spacing of 0.15 m to 1 m is used. For columns, above 1 meter in diameter, 0.5 m can be used as an initial estimate.

Column diameter is influenced by the vapor flow rate in the column. The trays can not have excess liquid entrainment or high pressure drops; therefore, vapor velocity in the column must be maintained at a reasonable level.

An equation based on the Souders and Brown equation can be used as an estimate for the max allowable superficial vapor velocity,

<math>\hat u_v = (-0.171l_t^2 + 0.27l_t - 0.047){\frac{\rho_L - \rho_v}{\rho_v}}^{1/2}</math>

where <math>l_t</math> is the plate spacing in meters, <math>\rho_L</math> is the density of the liquid stream, and <math>\rho_V</math> is the density of the vapor stream.

Column diameter, <math>D_c</math>, can then be estimated using the relation,

<math>D_c = \sqrt{\frac{4\hat{V_w}}{\pi\rho_v\hat{u_v}}}</math>

where <math>\hat{V_w}</math> is the maximum vapor rate in kg/s (Towler et al., 2013).

===Distillation Applications===

Distillation is a process that can be implemented in various scales. There is both laboratory scaled distillation as well as very large industrial distillation. Other applications for distillation include food/alcohol processing and herb distillation for the perfume and medical industries. Typically laboratory scaled distillation occurs in batches whereas industrial distillation (e.g. fractional distillation of crude oil) occurs continuous with a constant distillate and bottom effluent streams.

Some applications of distillation are concerned the top stream only, some the bottom stream only and others both streams can be used for future products. In alcohol distillation for example, the water that is separated from the ethanol/water binary solution is discarded as waste water. In fractional distillation of crude oils, the heavy hydrocarbons at the bottom of the column are collected and sold along with the light hydrocarbons that appear in higher side draws (Wankat, 2012).

===Example Case: Ideal Distillation===

Assume an equimolar mixture flowing at 10 mol/s of 20 mol% n-pentane, 30 mol% n-hexane, and 50 mol% n-heptane. Separate the mixture into 3 products: 99% pure n-pentane, 99% pure n-hexane, 99% n-heptane. Assume the feed and products are all liquids at the bubble points. There are two process alternatives to consider in this example. The direct sequence removes the most volatile species, pentane, in the first column, and then separates hexane and heptane in the second column. The indirect sequence separates the heaviest product, heptane, and then separates pentane from hexane in the second column. This example will consider the direct sequence. Next, we must decide if these species exhibit fairly ideal behavior during distillation. Since the n-alkanes have very similar properties, it is safe to assume they will display close to ideal behavior. The next step is to look up the boiling points of the 3 species. In this case, the normal boiling points of pentane, hexane, and heptane are 309 K, 342 K, and 372 K, respectively. Also, it is a good idea to look up relative volatilites, to further verify near-ideality of the mixture, but also to obtain the information necessary for the Underwood method, which we will employ to obtain a solution. The next step is to write out material balances based on molar flows and the design specifications. They go as follows:

<math>\mu_I(nC5) + \mu_{II}(nC5) = 2 mol/s</math>

<math>\mu_I(nC6) + \mu_{II}(nC6) + \mu_{III}(nC6) = 3 mol/s</math>

<math>\mu_{II}(nC7) + \mu_{III}(nC7) = 5 mol/s</math>

<math>\mu_I(nC5) = 99\mu_I(nC6)</math>

<math>\mu_{II}(nC5) = (5/990)\mu_{II}(nC6)</math>

<math>\mu_{II}(nC7) = (5/990)\mu_{II}(nC6)</math>

<math>\mu_{III}(nC7) = 99\mu_{III}(nC7)</math>

where <math>\mu</math> represents the molar flow, and the subscript represents the product stream.

Solving this system of equations:

<math>\mu_I(nC5) = 1.985\ mol/s</math>

<math>\mu_{II}(nC5) = 0.015\ mol/s</math>

<math>\mu_I(nC6) = 0.020\ mol/s</math>

<math>\mu_{II}(nC6) = 2.930\ mol/s</math>

<math>\mu_{III}(nC6) = 0.050\ mol/s</math>

<math>\mu_{II}(nC7) = 0.015\ mol/s</math>

<math>\mu_{III}(nC7) = 4.985\ mol/s</math>

At this point we have enough information to use Underwood's method to estimate the minimum vapor flows in the column. The following three equations are used in Underwood's method:

<math>\sum_i \frac{\alpha_{ik}}{\alpha_{ik}-\phi}f_i = (1-q)F</math>

<math>(R_{min}+1)D = \sum_i \frac{\alpha_{ik}}{\alpha_{ik}-\phi}d_i = V_{min}</math>

<math>\bar R_{min}B = -\sum_i \frac{\alpha_{ik}}{\alpha_{ik}-\phi}b_i = \bar V_{min}</math>

where <math>\alpha_{ik}</math> is the relative volatility of species <math>i</math> to species <math>k</math>, <math>f_i</math> the molar flow of species <math>i</math> in the feed, <math>q</math> the fraction of the feed that joins the liquid stream at the feed tray, <math>F</math> the total molar flow of the feed, <math>D</math> the molar flow of the distillate, <math>R_{min}</math> the minimum reflux ratio <math>(=L_{min}/D)</math>, <math>d_i</math> the molar flow of species <math>i</math> in the distillate, <math>V_{min}</math> the minimum vapor flow possible in the top section of the column to accomplish the desired separation, <math>\bar R_{min}</math> the minimum reboil ratio <math>(=\bar V_{min}/B)</math>, <math>b_i</math> the molar flow of species <math>i</math> in the bottoms product, and <math>\bar V_{min}</math> the minimum vapor flow in the bottom section of the column. The final variable, <math>\phi</math>, will be solved for using the first Underwood equation, and it's value will be decided based on the relative volatilities of the key components in the column.

So, after solving the first Underwood equation, we get two values for <math>\phi</math>, 3.806 and 1.462. Because 3.806 is between the relative volatilities of the key components, we will substitute that value for <math>\phi</math> into the second Underwood equation. Doing so for both columns gives <math>V_{min} = 6.4\ mol/s</math> for the first column and <math>V_{min} = 8.9\ mol/s</math> for the second column, for a total minimum vapor flow of 15.3 mol/s. The process would then be repeated for the indirect sequence, and the decision for which process to use would be justified by the process with the overall minimum vapor flow (Biegler et al., 1997).

==Absorption==
===Description of Absorption===
Another separation process used in industry is absorption, which is used to remove a solute from a gas stream. It accomplishes this by contacting the gas mixture with a liquid solvent that readily absorbs the undesirable components from the gas stream, purifying the gas stream. This separation process is determined by the inputs of the liquid flow rate, temperature, and pressure.

The absorption factor, which can be determined mathematically, determines how readily a component will absorb in the liquid phase. The absorption factor of component i is

<math>A_i=L/K_iV</math>

where <math>L</math> is the liquid flow rate entering the column, <math>V</math> is the vapor flow rate entering the column, and <math>K_i</math> is the vapor/liquid equilibrium ratio for component i (Peters & Timmerhaus, 2003). Higher absorption factors result in higher absorptivity into the liquid and a decrease in the number of trays required for separation, however a diminishing return occurs after the absorption factor is greater than 2.0. An absorption factor of 1.4 is most commonly used.

In general absorption can be seperated into two overarching categories, physical and chemical absorption. In physical absorption, the unwanted solute in the gas is absorbed into the liquid phase because solubility of the component is higher in the liquid phase than the gas phase. In chemical absorption the solute is removed from the gas via a reaction with the solvent, this reacted product is then transported into the liquid phase (Danckwerts 1965). There are two types of chemical absorption reversible and irreversible. Generally reversible chemical absorption is preferred as the solvent can be put through a stripper and regenerated so it can be recycled back to the absorption process (Wankat, 2012).

===Absorption Apparatus===

There are five major apparatus used for absorption in industrial application. These five pieces of equipment are spray absorbers (or towers), ejector (venturi) scrubbers, packed columns, trayed columns, and film absorbers (Schmidt, 2012).

==== Spray Tower vs Ejector Scrubber ====

In both '''spray tower''' and the '''ejector scrubber''' nozzles are employed to produce small solvent droplets. These small droplets increase the surface area of the liquid to gas contact allowing for the maximum amount of mass transfer to occur between the gas mixture and the liquid. The major difference between the two nozzle equipment designs is the configuration and type of nozzles. In the ejector scrubber shown in Figure 3 there is a single nozzle that is generally a higher pressure spray nozzle that produces finer solvent drops allowing for an even greater amount of mass transfer enabling better physical absorption (Schmidt, 2012).
[[File:Ejectorventuri.jpg|thumb|200px|center|Figure 3. Ejector Scrubber (US EPA, 2006)]]
'''Spray towers''' on the other hand generally have many nozzle at different heights where the liquid solvent will be sprayed out of to contact the gas running through the tower. This design is used in order to ensure the gas contacts the liquid as throughout the tower. These nozzles are lower pressure than a ejector scrubbers nozzle and thus physical mixing is worse in this configuration. Since physical mixing is generally worse in this configuration it is usually used in conjunction with a chemical absorption process. The other major difference between the ejector scrubber and the spray tower is that gas and liquid flow is cocurrent in the former while it is countercurrent in a spray tower. A spray tower absorber is shown below in Figure 4 (Schmidt, 2012).
[[File:SparyTowerAbsorber.jpg|thumb|200px|center|Figure 4. Spray Tower Absorber (US EPA, 2006)]]

==== Tower Type Absorption Apparatus ====
'''Packed column absorbers''' and '''tray column absorbers''' have very high efficiencies for the removal of an unwanted solute in the gas stream. The major disadvantage a trayed column has when compared to a packed column is the pressure drop. The pressure drop in a packed column is generally very low, whereas in between each tray of a trayed column pressure drop can be quite large. However the advantages inherent to trayed columns become clear when one needs the solvent to have a high concentration of the component to be removed from the gas stream. This is most important in the case where there is a very low concentration of the component in the gas stream and the specification states the solvent must contain a high concentration of that component. In this case the flow rate of the solvent may not be high enough for a packed column, however in a trayed column the solvent flow rate can be near zero for operation (Schmidt, 2012). Packed and trayed column internals are very similar to the setups found in the respective distillation columns.

For a '''trayed column''' the plate efficiency can be calculated using O'Connell's Correlation which invovles the Henry's Law constant, total system pressure, and solvent viscosity at the operating temperature (Towler & Sinnott, 2013).

<math>x=0.062*\frac{\rho_s*P}{\mu_s*H*M_s}</math>

where
<math>x</math> is the tray efficiency,
<math>\rho_s</math> is the density of the solvent in <math>kg/m^3</math>,
<math>P</math> is the total pressure of the system in <math>N/m^2</math>,
<math>\mu_s</math> is the solvent's viscosity in <math>mNs/m^2</math>,
<math>H</math> is the Henry Law constant in <math>1/(Nm^2*(mol fraction))</math>,
and <math>M_s</math> is the molecular weight of the solvent.

A packed towers height can be determined using the equations below when concentration of solute is below 10% so that the assumption that the flow of gas and liquid will be essentially constant throughout the column holds (Towler & Sinnott, 2013). The height of packing <math>Z</math> is given by the following equation:

<math>Z=\frac{L_m}{K_G*a*P}*\int\limits_{y_2}^{y_1} \frac{dy}{y-y_e}\,</math>

where <math>P</math> is the total pressure, <math>a</math> is the interfacial surface area per unit volume, <math>y_1</math> and <math>y_2</math> are the mol fractions of the solute in the gas stream at the bottom and top of the column respectively, <math>G_m</math> is the molar gas flow rate per unit cross-sectional area, and <math>y_e</math> is the mole fraction of solute in the gas that would be in equilibrium with the liquid concentration.

The first half of the equation before the integral can be called the height of an overall gas-phase transfer unit <math>H_G</math> and the second part of the equation is the number of overall gas-phase transfer units or <math>N_G</math>. Using these definitions the above equation can be simplified to

<math>Z=H_G*N_G</math>

These equations assist in sizing an absorption column (Towler & Sinnott, 2013).

==== Film Absorber ====
The final absorber the film absorber is generally used in the case where the heat of absorption must be removed. The film absorber operates by sending the gas and solvent through a heat exchanger where the solvent creates a thin film on the walls of the tubes and the gas flows through the interior allowing for solute transfer. The good heat transfer present in a film absorber makes it preferable for situations where low temperatures are required for a high recovery of the solute (Schmidt 2012).

===Industrial Absorption Processes===
An industrial example is lean oil absorption, which is used to separate nitrogen and other impurities from natural gas. A lean oil is contacted with low quality natural gas, and the methane is selectively absorbed by the lean oil, leaving the impurities behind. The methane is subsequently regenerated from the rich oil as high quality natural gas (Petrogas Systems, 2014).

Other common industrial practices of absorption come from sour gas treatment. Amine gas treating is used to remove hydrogen sulfide or carbon dioxide from gas streams via a reversible chemical absorption. In amine gas treating the sour gas is fed to the bottom an absorber where amine solution is fed to the top along with any necessary make up water. The sour gas components are absorbed into the amine via a chemical absorption method. Sweet gas leaves the top of the absorber whereas the amine out of the bottom, now rich with acidic components is sent to a regenerator where the acid gas components are stripped and the acid gas is generally sent to a flare whereas the amine now lean again is recycled back into the first absorber (Miller & Zawacki, 1978). Figure 5 below shows the typical setup of an amine plant. Another type of sour gas treatment that uses absorption is Merichems LO-CAT process which uses a chelated iron to remove hydrogen sulfide from feed gas in the absorption column (Merichem 2015).
[[File:AmineTreating.png|thumb|400px|center|Figure 5. Amine Gas Treating Plant Schematic]]

==Stripping==
This process separates solutes from solvents (often after absorption, to purify the solvent so that it can be recycled to an absorber). Stripping will depend on the vapor and liquid flow rates, as well as the temperature and pressure of the column. There is a temperature drop down the column, so columns generally have either an increased operating temperature or decreased operating pressure.

The stripping factor of component i is

<math>S_i=K_iV/L</math>

where <math>K_i</math> is the vapor/liquid equilibrium ratio, <math>V</math> is the vapor flow rate entering the column, and <math>L</math> is the liquid flow rate entering the column, will determine how much of solute i will be stripped from the liquid into the vapor phase (Peters & Timmerhaus, 2003). The usual range for the stripping factor is between 1.2 and 2.0, with a stripping factor of 1.4 being most economic.

An example of stripping in industry is the deodorization of food items such as oils. The oil is heated and allowed to trickle down the column while steam flows up from the bottom of the column. At the vapor-liquid interface, volatile components of the oil transfer to the steam and are carried off the top of the column, leaving a purified oil product (Alfa Laval, 2014).

==Bioseparations==
===Importance===
As our ability to manipulate and engineer biological systems improves, biological products are becoming an increasingly important source of therapeutics and fuels. The production of fuels from biomass via either the enzymatic breakdown of a feedstock or the secretion of usable lipids from algae is a promising new energy source. Additionally, enzymes, antibodies and other therapeutic proteins have been applied to the treatment of a wide range of diseases. Although each process requires its own set of separations, all follow the same basic format: separation of biomass, product isolation, and product purification (Belter et al., 1998). This section will provide examples of unit operations in each step. Ultimately, the choice of separation process and unit operations will depend on the specific process and product. The descriptions below are examples of the most common bioseparation operations within the general platform (Harrison et al., 2003).

Bioprocesses begin with fermentations or growth operations. In biofuel production processes, this may involve growing algae or breaking down corn or cellulosic biomass. For the production of therapeutics, mammalian or bacterial cells may be grown in a fermentor and the product secreted into the supernatant or harvested from the cells.

===Biomass Separations===
After fermentation and product production, the solid biomass must first be separated from the desired product. If the product is secreted from the cells, this can be done immediately after fermentation ends. If the product is not secreted, the cells must first be lysed.
Cell lysis is the process of lysing, or breaking, the cell in open. Mechanical lysis is the simplest, and involves physically breaking the cell either by mashing (think mortar and pestle) or blending the cells into a homogenous solution in a homogenizer. Chemical lysis is another method, achieved by introducing an osmotic shock or chemically degrading the cell membrane. Additional separation can be achieved by flocculation, which is the process of aggregating biomaterial by charge neutralization or bridging. These larger complexes are easier to separate from smaller molecules (Harrison et al., 2003).

The next step is removing the unwanted biomass from the product in solution. Separation by centrifugation or sedimentation are the most common, although filtration is sometimes also used for processes where a biomass cake is desired. Both methods utilize density differences to separate the product from the solid biomass (Towler and Sinnott, 2013).

====Sedimentation====

Sedimentation relies purely on the force of gravity, while centrifugation speeds the settling process by subjecting the cells to a centrifugal force. Sedimentation in a settling tank is the simplest method of solid-liquid bioseparation. In this process, biomass in a tank is simply allowed to settle to the bottom over time. While this process is inexpensive, requires little energy and can separate out large volumes of biomass, it generally requires long time periods and is only mostly in very large-scale processes where active centrifugation is difficult (Belter et al., 1998).

====Centrifugation====
Centrifuges are widely utilized across many processes, and thus a wide variety of scales and designs have been developed. Disk-stack centrifuges, in which the solid phase is deposited onto “shelves” in the center of the spinner and liquid phase is pushed to the outside, are some of the most commonly used centrifuges in industry. They are especially suited to biomass separation processes because they can be built on a large scale and are ideal for separating fine solids from liquids. [[File: Disk_stack_centrifuge_towler.png|frame|center|Fig. 6: Diagram of a disk-stack centrifuge (Tolwer et al, 1997).]] Tubular bowl centrifuges are also common and can reach separation efficiencies of up to 90%. Heavier products accumulate along the sides of the bowl, while the light phase flows out the top. They separate products by can be used both to separate solids from liquids and immiscible liquids, such as and oil product and an aqueous broth (Tolwer and Sinnott, 2013). [[File: tubular bowl centrifuge towler.png|frame|center|Fig. 7: Diagram of a tubular bowl centrifuge centrifuge (Tolwer and Sinnott, 2013).]]

Centrifugation scale-up is made easier by sigma analysis, which allows for the estimation of appropriate feed rates for different size centrifuges. The sigma factor is dependent on the inner and outer radius of the centrifuge, the angular velocity, and the sedimentation velocity of the solid particles being separated. It can be thought of as the characteristic cross-sectional area with units of [length]2. The sedimentation velocity can be calculated by

<math>v_g={\frac{2a^2(\rho-\rho_0)}{9\mu}}g</math>

where <math>v_g</math> is the sedimentation velocity, <math>a</math> is the cell or biomass particle diameter, <math>\rho</math> is the particle density, <math>\rho_0</math> is the fluid density, and <math>\mu</math> is the fluid viscosity. The volumetric flow <math>Q</math> can be estimated by

<math>Q=(v_g)(\Sigma)</math>.

The equality

<math>{\frac{\Sigma_1}{\Sigma_2}}={\frac{Q_1}{Q_2}}</math>

can be an easy way to estimate equivalent flow rates between a small-scale centrifuge 1 and larger centrifuge 2 (Harrison et al., 2003).

====Example: Centrifugation Scale-up====

You are trying to separate a cell of radius 0.4 <math>\mu</math>m with a density of 1.05 g/cm3 from broth of mostly water (density of 1 g/cm3 and viscosity of 0.01 g/cm s). The sigma factor of the centrifuge you are using is 1 x 106 cm2. A] What volumetric flow rate should you use? B] If you want to scale up the process to a centrifuge with <math>\Sigma</math> = 3 x 106 cm2, what flow rate would you use in the larger centrifuge?

Solution:
A] Using the equation for <math>v_g</math>, and being mindful of units, the sedimentation velocity equals 1.74 x 106 cm/s. The flow rate, then, equals

<math>Q=(1.74 x 10^-6)(1,000,000) = 1.74 cm^3/s = 0.104 L/min</math>.

B] Keeping in mind that for the same process, <math>v_g1 = v_g2,</math> and rearranging the sigma factor equality, the new flow rate is

<math>Q_2 = {\frac{\Sigma_2 x Q_1}{\Sigma_1}} = {\frac{(3 x 10^6)(0.104)}{1 x 10^6}} = 0.313 L/min </math>

===Product Isolation===
Liquid-liquid separation, to extract the product from the aqueous phase, is much less straightforward than liquid-solid extraction. Many methods - especially adsorption, filtration, and precipitation - are similar in principle to operations found in other, non-biological separations. The exact separations used depend on the nature of the product and the scale of the process. These processes are nearly identical to their non-biological counterparts, and their description is left to other sections.

Particular care needs to be taken with protein products because of their instability, and the selection of an appropriate solvent or adsorbent is crucial to a successful process (Harrison et al., 2003).

===Product Purification===
The final steps of protein purification and polishing remove any remaining contaminants and bring the concentration of product to an appropriate value for applications. Purification processes for food-grade and medical products can be extensive, as sterility and high purity are essential. Purification in fuel-producing processes may be less extensive, depending on the process. Chromatography and crystallization are two common steps in purification and are especially used in industrial scale protein production. Several different types of chromatography exist with the ability to carry out different types of separations.
Chromatography is similar to adsorption in that it relies on differences in affinity between solutes and a solid surface. A solution is eluted through a column containing a solid resin with various affinities for the substances in solution. In adsorption, the solutes are evenly saturated throughout the column. Chromatography differs in that solutes are deposited a resin phase before the column is flushed with an elution solvent specific that results in solutes eluted in bands.

==== Ion Exchange Chromatography ====

There are two main types of ion exchange columns—anion and cation. Anion exchange resins have a positive charge and are used to retain products with a negative charge. Cation exchange resins have a negative charge and are used to retain products with a positive charge. The pH of the elution buffer is change to force a specific solute to wash out, depending on whether the pH of the buffer is above or below the isoelectric point of the solute (Belter et al., 1998). This is especially useful for the separation of protein product (including antibodies), nucleic acids, and other charged molecules. When the solutes have sufficiently different isoelectric points, the pH of the buffer is manipulated to affect the solute charge and force the product to elute while the solute remains preferentially bound to the resin, or vice versa (Harrison et al., 2003). In general, the most strongly charged molecules will remain in the column for a longer period of time. Elution washes through the weakly bound ions before the more strongly bound ions. Different speeds of elution can be visualized as in figure 8.

[[File:chromatography.png|frame|center|Fig. 8: Illustration of product bands in an elution chromatography column (Belter et al., 1998).]]

==== Size Exclusion Chromatography ====

In gel filtration chromatography, small molecules are "trapped' by the porous resin and take longer to flow through the column. Larger products will elute first because the smaller molecules are better able to penetrate the resin. This forces them to take a much longer path through the column, which means it takes longer for them to elute. This operation is often used when there is a distinct difference in size between the desired product and other solutes.

==== Affinity Separations ====

Affinity chromatography is very similar to ion exchange chromatography in that the interactions between the material in the column and the molecules in the feed. The main difference is that affinity chromatography can rely on a great variety of types of interactions. Two very common types of affinity are exploited in affinity chromatography columns. The first is immunoaffinity. Proteins are specifically bound by antibodies which can be incorporated onto beads and used in chromatography. Antibodies are designed to bind only a single protein, so these interactions are considered to be highly specific. The protein can be eluted using a buffer that changes the pH or salinity in the column, which adversely affects binding (Hage, 1999).

Proteins can be separated from very complex and unrefined mixtures using this method because of the great specificity. Antibody-protein interactions are so specific that individual proteins can be isolated from blood samples as shown in the figure below. The main drawback of this method is simply that specific interactions between proteins and binding targets do not always exist. Antibodies exist for many proteins, but for others they must be customized or created using some sort of evolution process. This is not a trivial task and can require a significant capital investment (Hage, 1999).

[[File:Affinity_Chromatography_Example.PNG|frame|center|Fig. 9: Affinity purification separating fibrinogen from human plasma using an anti-fibrinogen antibody (Hage, 1999).]]

The other main type of affinity chromatography is based on protein specific tags and the molecules or surfaces to which they bind. One of the most common types of protein tags used is the polyhistidine tag. This tag consists of 6-8 consecutive histidine residues which can be added to the exterior of the desired protein product. The addition of this tag requires alterations to the coding sequence of the protein. The polyhistidine tag binds strongly with nickel and cobalt ions. The product with the tag can then be eluted with imidazole—a small molecule with the same structure as the functional group of the amino acid histidine. Imidazole will bind the cobalt and nickel ions more strongly than the histidine in the tag. Along with chromatography, protein tag interactions can be leveraged with the use of beads that can be deposited directly into the solution containing the protein of interest (Lichty et al., 2005).

Different proteins can be purified in this manner with varying levels of efficiency. There is a very high dependence on the size of the protein, electronic properties, and steric considerations. The figure below demonstrates varying separation efficiencies with proteins of different sizes (HisPur Cobalt Resin - Thermo Fisher Scientific, 2016).

[[File:HisTagPurification.PNG|frame|center|Fig. 10: Protein gel demonstrating the separation of proteins from a sample. The far left column represents the most efficient separating conditions using cobalt and the far right column represents a negative control with no purification performed (HisPur Cobalt Resin - Thermo Fisher Scientific, 2016).]]

Several other types of tag-bead interactions can be utilized in separations processes. Maltose Binding Protein is a small protein that can be added to a protein of interest. It binds strongly with beads coated in immobilized maltose and can be released by flushing with maltose. As MBP is a full sized protein that typically must be removed from the protein of interest in order for it to be used. In this case, the site specific TEV protease is often used cleave MBP from the protein of interest. In addition, under specific circumstances, other unique tags can be used and provide varying levels of specificity in separations. The Flag tag, 3x Flag tag, Glut tag, and Strep tag. While these are all commonly used, the polyhistidine tag is the most popular because it gives the highest level of specificity. Some of the main purification tags are compared in the table below (Lichty et al., 2005).

{| class="wikitable"
|-
! Tag
! Size (aa)
! Resin
! Eluting Agent
! Capacity (mg/ml)
! Cost/10 mg
|-
| His
| 6
| Talon
| Imidazole
| 5-14
| $18
|-
| Maltose Binding Protein
| 396
| Amylose
| Maltose
| 3
| $12
|-
| Glutathione S-Transferase
| 218
| GSH-Sepharose
| Glutathione
| 10
| $36
|-
| Strep II
| 8
| Strep-Tactin Sepharose
| Desthiobiotin
| 0.1
| $293
|-
| FLAG
| 8
| Anti-FLAG M2 MAb Agarose
| Flag peptide
| 0.6
| $1045
|}

Unfortunately, based on the contents of the solution, the characteristics of the protein, and the specific product being used, there are extremely large amounts of variation in yield and purity for each type of tag. Therefore, it is crucial to allocate time towards preliminary tests and scale-up experiments with the exact solution from which the protein of interest is being separated. The his tag is typically the most used because, as seen in the table above, it has one of the lowest costs and highest yields. In addition, because it is only a 6 amino acids, it often is not cleaved from the protein of interest which minimizes the number of steps required in the overall purification process (Lichty et al., 2005).

==== Crystallization ====

Crystallization, or the formation of solute crystals from a solution, is especially useful in biomolecule separations because it is possible to obtain a 99.9%+ product purity. In crystallization, a diluent is added to the homogeneous solution that reduces the solubility of the product to the point that it “falls out” of solution and crystallizes. It is similar to precipitation but results in the formation of crystals rather than unordered aggregates. Crystallization can be used on a laboratory scale for determining protein structure, on on the industrial scale for antibody and therapeutic protein productions. Batch crystallizers are often used in industry because of their simplicity and inexpensiveness compared to continuous crystallization (Harrison et al., 2003).

==Membrane Separation==
Membrane separation takes advantage of the selective permeability of membranes; they allow certain particles to pass through and selectively stop other, generally unwanted, particles. The component that passes through is called the permeate and the component stream that is rejected is called the retentate or concentrate. The applicability of membranes comes from the fact that their selectivity is determined by their pore size, which can be controlled during the creation of the membranes. Additionally, Membrane processes do not require heat meaning they generally require less energy than conventional separations technology such as distillation and crystallization. Membrane separations processes are generally classified as microfiltration, ultrafiltration, or nanofiltration depending on the size of the particles to be filtered out.

[[File:membrane techs.png|frame|Figure. Cutoffs for different membrane categories]]

===Membrane Selection, Construction, and Flow Geometries===
Membrane permeability and selectivity are the two most important factors to consider when selecting a membrane. For gas separations, the permeation of the gas is usually facilitated by the gas dissolving in the membrane on one side and then evaporating on the permeate side. Therefore permeability depend largely on the solubility of components in the membrane.

The two most commonly utilized membrane configurations are hollow fiber and spiral wound. Hollow fiber is generally the most commonly utilized module for gas separations. These are formed by gluing the two ends of the hollow fiber to a resin forming a closure. The fibers are housed in a shell much like a heat exchanger. The feed flows past thousands of tubes with the permeate flowing into the hollow tubes and out the closure. The retentate then flows out of the shell not having gotten through the membrane.

[[File:hollow membrane.jpg|frame|Figure. Hollow fiber membrane module]]

Spiral wound membranes are created by sealing two membrane sheets back to back on three edges to form a sort of pocket. This fourth open edge is then attached to a porous tube which allows permeate to go through it. Several membrane pockets are attached to a single tube and wrapped around in a spiral.

[[File:spiral membrane.jpg|frame|Figure. Spiral wound membrane module]]

Flow geometry is usually either dead-end geometry or cross flow geometry. In dead end, the fluid flow is normal to the membrane surface while cross flow is parallel to the membrane surface. Dead end geometry is usually used with hollow fiber membranes while cross flow is used with spiral wound membranes. Each geometry has advantages and disadvantages. Dead end geometry is generally cheaper to set up and therefore has lower initial capital costs. However, it is very vulnerable to membrane fouling, which reduces the effectiveness of the membrane. This is usually the geometry set up for small scale lab experiments. The tangential flow devices are more cost and labor-intensive, but they are less susceptible to fouling due to the sweeping effects and high shear rates of the passing flow. Most commercial industrial membrane separations are done using spiral wound cross flow membrane modules.

The first big breakthrough for membrane technology was the creation of the asymmetric membrane. In this type of membrane, there is a very thin dense polymer portion of the membrane that is cross linked with less dense, spongier, support polymer. The dense part of the membrane does most of the filtering while the spongy section provides structural support ensuring the membrane will not break. This allows the dense portion to be thinner which increases overall flux through the membrane. The early asymmetric membranes were made with all the same polymer. State of the art membranes are called thin film composite membranes. They are essentially asymmetric membranes were the dense thin portion is made from a different polymer than the spongy support. This enables the best support polymer to be paired with the best filtering polymer allowing for much more effective membranes. Current thin film composite membranes are usually made with a polyamide thin film layer supported by polysulfone.

===Designing a Membrane System===

When designing a filtration system you must first specify your permeate (product) flow rate and the flux of your membrane along with the membrane area of your selected unit. Number of units can be found using the following equation

<math>N_{E} = \tfrac{Q_{p}}{f * S_{E}} </math>

Where Q_p is the permeate flow rate, f is the flux, and S_E is the area of the selected membrane.

Once you have also specified a recovery, you can choose a number of stages (series membranes as opposed to parallel). Generally if the desired recovery is below 50%, the separation can be done in one stage. Recoveries higher than this are generally best done with 2 stages. More stages can be used to achieve very high recoveries but these are rare. Once you have chosen a number of stages, you can use the following equations to determine the number of membranes in each stage

<math>\begin{align}
R & = {\tfrac{1}{1 - Y}}^{1/n} \qquad & R &= \tfrac{N_{E_{i}}}{N_{E_{i+1}}}
\end{align}</math>

Where Y is the recovery and n is the number of stages and R is the staging ratio. N_E_i is the number of membrane elements in stage i.

===Applications===
====Food Industry====
Due to the fact that MD can be conducted at relatively low feed temperatures, it was successfully tested in many areas where high temperature applications lead to degradation of the process fluids especially in food processing. It was demonstrated that MD can be used for the concentration of milk, for the recovery of volatile aroma compounds from black currant juice, and for the concentration of many other types of juices including orange juice, mandarin juice, apple juice, sugarcane juice, etc.
====Reverse Osmosis====
Reverse osmosis is the most widely used membrane separation process. In this process, fresh water passes through the membrane while dissolved salts and other solids are rejected and stay in the concentrate. In this process, feed water is pressurized in order to overcome the osmotic potential difference between the salty retentate and the fresh water desired. These processes are generally run using spiral wound membrane cylinders using a cross flow setup.

===Membrane Model===
The two most important components when considering different membranes are the permeability, which will determine flux through the membrane, and selectivity, which will determine what passes through the membrane and how much. The flux through a membrane is defined as:

<math>M_i = \tfrac{P_i}{\delta}(p_{\text{i,f}} - p_{\text{i,p}})</math>

Where <math>M_i</math> is the molar flux of component i, Pi is the permeability of the membrane for component i, δ is the membrane thickness, and <math>p_(i,f)</math> and <math>p_(i,p) </math> are the partial pressures of component i on the feed side and permeate side respectively.
The average flux across a long cylindrical membrane such as the spiral wound module is given by:

<math> \int_0^Lm \frac{M_i,dx}{L_m}</math>

Where Lm is the length of the cylinder and x is length in meters

Membrane selectivity of the ideal separation factor is given as the ratio of the permeability of one substance over another as shown:

<math> S_{ij) = P_i/P_j </math>

where <math>S_{ij} </math> is the selectivity of the membrane for component i over j.

==Cyclones==
Centrifugal Separators, more commonly know as Cyclones, are one of the most widely used gas-solid separators. They are typically used for particles that are 5μm or larger, but if agglomeration occurs they can sometimes separate particles as small as particles as small as .5μm. These cyclones can be designed for high efficiency separations or for high throughput (Towler, 2013).

[[File:cyclone.png|thumb|center|Reverse-flow Cyclone (Towler, 2013)|500x300px]]

Cyclones are popular in industry because of their low capital costs, low maintenance costs, low pressure drop, temperature and pressure limitations are only caused the materials, they work as dry equipment, and they don't take up much space (EPA).

They are not without their disadvantages. They have low collection efficiencies for small particles, they are unable to handle sticky materials and high efficiency units often have large pressure drops (EPA).

===Design===
Design Procedure
There are multiple design methods for cyclones. The deign method discussed below is the Stairmand method. This involved creating two designs one of which was for high throughput and the other which was designed for high efficiency. Using these base designs future cyclones can simply be estimated using scaling factors (Towler, 2013).

1. Select whether efficiency or performance is more important for your design

2. Determine the particle size distribution in your stream

3. Estimate the number of cyclones needed

3. Calculate the cyclone diameter

4. Calculate the scale up factor

5. Calculate the cyclone performance and overall efficiency

6. Calculate the cyclone pressure drop

7. Cost the system

====Particle Size====
The diameter of particles separated by cyclones are governed by this design equation. This equation is dependent upon the standards of the cyclone and how much you are trying to deviate from its standard operating conditions. Looking into the equation it shows that a larger internal cyclone diameter, higher flowrate, higher difference in density, and lower viscosity result in smaller particles being seperated. The whole term after d1 is known as a scale up factor. When designing a cyclone you would determine the flow rate, viscosity, particle size, and density difference would be known prior to using this design equation, so this would be used to optimize the diameter of the cyclone you were designing (Towler, 2012).

[[File:Towler 10.8.PNG|center|Reverse-flow Cyclone (Towler, 2013)|500x300px]]

d1=mean diameter of particle separated at the standard conditions, at the chosen
separating efficiency

d2= mean diameter of the particle separated in the proposed design, at the same
separating efficiency

Dc1 =diameter of the standard cyclone

Dc2 =diameter of proposed cyclone

Q1 =standard flow rate

Q2 =proposed flow rate

Δρ1 =solid-fluid density difference in standard conditions

Δρ2 =density difference, proposed design

μ 1=test fluid viscosity

μ 2=viscosity, proposed fluid.

[[File:Towler10.44.PNG||thumb|center|(a) Standard Cyclone Dimensions (b) High Gas Rate Cyclone (Towler, 2013)|500x300px]]

====Efficiency====
Efficiency is the percentage of total solids that are removed from the stream The graphs below were experimentally for the two standard designs Stairmand Method. This graph has the efficiency vs the particle size (Towler, 2012).

[[File:Towler10.45.PNG||thumb|center|Performance curves, standard conditions. (a) High-efficiency cyclone performance curves, standard conditions. (b) High gas-rate cyclone (Towler, 2013)|600x400px]]

If you have to scale the efficiency it can be estimated by the point where a horizontal line from the efficiency of the original particle size intersects the new particle size. A demonstration is on the graph below.

[[File:Towler 10.46.PNG||thumb|center|Scaled Performance Curve(Towler, 2013)|500x300px]]

====Pressure Drop====
One of the most important design factors in a cyclone is its pressure drop. The pressure drop occurs because entry and exit losses, kinetic energy losses, and friction losses that occur inside the cyclone. The pressure drop can be calculated using the following equation (Towler, 2013).

[[File:Towler 10.9.PNG|center|Reverse-flow Cyclone (Towler, 2013)|500x300px]]

The variables are defined as follows:
DP = cyclone pressure drop, millibars;
rf = gas density, kg/m3
;
u1 = inlet duct velocity

u2 = exit duct velocity

rt = radius of circle to which the center line of the inlet is tangential

re = radius of exit pipe

φ =factor from figure below

ψ = fc(As/A1)

fc = friction factor

As = surface area of cyclone exposed to the spinning fluid

A1 = area of inlet duct

[[File:Towler_10.47.PNG|thumb|center|Cyclone Pressure Drop Factor (Towler, 2013)|600x400px]]

====Cost Estimation====
The following values are represented in 2002 dollars. Cost is mostly a function of the amount of gas that need to be passed through the cyclone and the units are in standard meters cubed per second (sm3/sec).

The costs can be broken down into three categories:

Capital Costs: $4600-$7600 sm3/sec

Operations and Maintenance: $1500-$18,000 sm3/sec

Annualized Cost: $2800-$29000 sm3/sec

This all comes to $.47-$440 per metric ton of pollutant removed. The high variation in these costs have to do with smaller units being much less efficient per unit of pollutant removed. Furthermore the cost is also effected by the amount of particulates that are in the gas to begin with. For example a gas with 10 g/sm3 and air with 100sm g/sm3 would have similar pumping costs, but very different cost per unit of solids removed (EPA).

==Other Separation Processes==
===Extraction===
Liquid-liquid extraction is a process for components with overlapping boiling points and azeotropes. The process requires a solvent such that some of the components of the mixture are soluble, and then the components will be separated based on this solubility in the liquid. This process can operate at moderate temperatures and pressures, so is not very energy intensive. However, a distillation column is required to extract the solvent for recycle. More recently, supercritical fluids have replaced liquid solvents in some processes for L/L extraction, due to the solute’s ability to more rapidly diffuse through them. The issue with these fluids, however, is that they must be operated at extremely high pressures and temperatures, increasing both capital and operating expenses of the process (Peters & Timmerhaus, 2003).

===Crystallization===
This process recovers solutes that have been dissolved in solution. The resulting product is in the solid phase. Depending on the material properties of the solute and solvent, the solute is recovered by precipitation after cooling, removal of solvent, or adding precipitating agents. Crystallizers are designed based on phase equilibria, solubilities, rates and amounts of nuclei generated, and rates of crystal growth. Every crystallization process is a unique system, so plant evaluation is usually required before complete implementation. Crystallization can be performed in both batch and continuous processes, and design features can control crystal size to an extent (Peters & Timmerhaus, 2003).

===Membrane Separation===
This separations process uses selectively permeable membranes to separate components in a mixture. Typically, one of the components will freely pass through the barrier while the other components will not. The stream that passes through the membrane is the permeate and the stream that does not pass is the retentate. The driving force behind this separation is a pressure gradient. Membrane separation is beneficial because it can separate mixtures at the molecular and small particle level. Furthermore, there is no phase change required so the energy input is low. Limitations of this process include achieving high product purity, incompatibility with certain stream components, low operating temperature, and low flow rates. Although membrane separation is generally not scaled up, examples of scaled-up membrane separation include seawater desalination and hydrogen recovery (Peters & Timmerhaus, 2003).

===Adsorption===
Adsorption involves an adsorbent and adsorbate. The adsorbent is typically a solid, and will typically separate the adsorbate from the stream. This process usually includes a desorption step that regenerates the adsorbent for further use. Raising the temperature or increasing the concentration of the adsorbate can reverse the adsorption process. Although the recycle of the adsorbent is a very economic design feature, the downside of this step is that it results in a cyclic process, which introduces complexity to the overall process. Industrial applications of this process are for bulk separations and gas purification. The adsorption/desorption process in these situations involves a large amount of heat transfer, which design engineers must take into account when sizing and selecting equipment material (Peters & Timmerhaus, 2003).

===External Field/Gradient Separation===
These separations use external force fields or temperature gradients to separate responsive molecules or ions. The use of these processes is fairly limited to a few specialized industrial applications (Peters & Timmerhaus, 2003).

===Settling and Sedimentation===
In settling processes, solid particles or liquid drops are separated from a stream by gravity. The stream can be in either the liquid or gas phase. For vapor-liquid mixtures, flash drums are generally used to separate the mixture. The velocity of the vapor must be less than the settling velocity of the liquid drops for this separation to occur. For liquid-liquid separation, the horizontal velocity of the fluid must be low enough to allow the low-density droplets to rise to the interface and the high-density droplets to move away from the interface and coalesce. In sedimentation, the result of the process is a more concentrated slurry. Typically a flocculating agent is used to aid in the settling process. One way to perform this separation is to use a cone-shaped tank with a slowly revolving rake that scrapes and moves the thickened slurry to the center of the cone for removal (Peters & Timmerhaus, 2003).

[[File:Example.jpg]]

====Clarifiers====
[[File:Circular_Clarifier.png|300px|thumb|bottom|Figure 9: Circular clarifier with some components labelled.]] [[File:Rectangular_Clarifier.png|300px|thumb|bottom|Figure 10: Rectangular clarifier with some components labelled.]]

Clarifiers are one of the methods used for the continuous removal of particulate solids from liquids through sedimentation by gravity. Applications include process water pretreatment, waste water treatment, and drinking water purification. Historically, clarifiers were originally developed to limit nutrient input into surface water due to fear of eutrophication. Today, they have a number of uses, particularly in wastewater treatment processes, metal removal, disinfection, and membrane pretreatment. The process helps removed dissolved solids, silt, and undesirable metals from the water, making it more suitable for downstream processes as well as human consumption (Wilson, 2005).

Clarifiers are typically used in conjunction with coagulation or flocculation agents, which promote dissolved particles to join into clumps and settle out of solution (Towler and Sinnot, 2012). Clarifiers typically consist of a large circular tank with a rotating rake at the base which scrapes settled solids towards the center. In the case of a rectangular clarifier, they are scraped to one side. Diagrams of both are represented in figures 9 and 10, respectively (NMED Surface Water Quality Bureau, 2015). Separated solids are allowed to settle to the bottom of the tank as a sludge, whereupon they are collected by the rake and disposed of properly. In the case of floating contaminants, it is possible for the clarifier to include a skimmer as well.

Clarifier efficiency varies with certain factors, including the settling characteristics of solids removed and the surface overflow rate of the tank. Clarifier efficiency can be found using the following relation:

<math>\begin{align}
E_{TSS} &= E_{TSSmax}\left ( 1 - e^\frac{\lambda}{SOR} \right )
\end{align}</math>

where <math>E_{TSS}</math> is the efficiency of total suspended solids (TSS) removal, <math>E_{TSSmax}</math> is the maximum possible efficiency, <math>\lambda \left [\frac{m}{d} \right ]</math> is the settling constant, and <math>SOR \left [\frac{m^3}{m^2 d} \right ]</math> is the surface overflow rate. The effect of flocculation chemicals on TSS can be seen in figure 11. However, it should be noted that chemical addition will increase sludge quantity and may have an adverse effect on plant aesthetics, which increases maintenance costs (Wilson, 2005).

[[File:Chem_Addition.png|200px|thumb|bottom|Figure 11: The effect of flocculating agents on total suspended solids removal in clarifiers.]]

=====Lamella Clarifiers=====

Lamella clarifiers use inclined plates in order to maximize the settling area for solids. Solids continue to settle into a hopper at the bottom of the tank while clarified water exits up through the inclined plates. This allows for the design of a smaller tank, which leads to large savings in capital costs. A lamella clarifier is pictured in figure 12.

[[File:Lamella_Clarifier.png|300px|thumb|bottom|Figure 12: A lamella clarifier with components labeled.]]

Typically, inclined plates are installed at an angle of 45 to 60 degrees and spaced 40 to 120 mm apart, which increases effective settling surface area by a factor of 6 to 12 compared to traditional clarifiers. For effective use, it is recommended that the Reynolds number be below 2000, Froude number higher than 10-5,and detention time be longer than 3 to 5 minutes. For this implementation, the equations are as follows:

<math>\begin{align}
N_{Re} &= \frac{VR}{\nu} \\
N_{Fr} &= \frac{V^2}{Rg}
\end{align}</math>

where <math>R</math> refers to the hydraulic radius, which is the cross-sectional area of the lamella, <math>V</math> is the liquid velocity, <math>\nu</math> is the kinematic viscosity, and <math>g</math> is the gravitational constant (Wilson, 2005).

=====Advantages=====

Clarifiers offer a proven, relatively inexpensive solution for solids removal. The chemical coagulants used are cheap and provide a low operating cost as well as simple maintenance. Construction is typically simple, leading to low capital costs and equipment that is easy to accommodate and maintain. Their design is also flexible, with various options such as skimmers and scrapers offering increased removal efficiency (Wilson, 2005). Operation of clarifier tanks also has lower energy requirements than membrane filtration for solids removal, given that most of the separation is aided by gravity. Water exiting clarifier units has a silt density index (SDI) averaging 4.0, which is low enough for further membrane treatment such as reverse osmosis (Prihasto, 2009).

=====Disadvantages=====

Clarifiers necessitate low turbulence to prevent resuspension of solids. This essentially requires a low entrance velocity, which can limit the production rate of certain processes or call for more clarifier units, which would drive up costs. Furthermore, clarifiers require frequent cleaning before sludge becomes too difficult to remove and reduces effectiveness. In the case of lamella clarifiers, sludge buildup on the inclined plates results in uneven flow distribution which could harm efficiency (US EPA, 2003). For this reason, maintenance requirements for lamella clarifiers are higher, but they can be reduced through the implementation of removable plates (Wilson, 2005). Clarifiers also only remove solids, so pH will not be affected, leading to the need for further pH adjustment (NMED Surface Water Quality Bureau, 2015).

=====Clarifier Design Calculations and Typical Design Values=====

======Detention Time======

Detention time (DT) is the time is takes for a unit of water to travel from the inlet of the clarifier unit to the outlet. During typical operations, the design value for this is 2 to 3 hours.

<math>\begin{align}
DT &= \frac{Tank\ Volume}{Influent\ Rate}
\end{align}</math>

======Surface Overflow Rate======

Surface overflow rate (SOR) measures the flow into the clarifier per square foot of surface area. Typical design values are 400 to 800 gal/day/sq. ft.

<math>\begin{align}
SOR &= \frac{Volumetric\ Flow\ Rate}{Surface\ Area}
\end{align}</math>

======Weir Overflow Rate======

Weir overflow rate (WOR) describes the flow in gallons per day per linear foot of weir. Typical values are 10,000 gal/day/ft.

<math>\begin{align}
WOR &= \frac{Volumetric\ Flow\ Rate}{Weir\ Length}
\end{align}</math>

======Solids Loading Rate======

Solids loading rate (SLR) describes the mass of solids in the clarifier influent per square foot of surface area. This value should not exceed 30 lbs/day/sq. ft.

<math>\begin{align}
SLR &= \frac{Solids\ Mass\ Flow\ Rate}{Surface\ Area}
\end{align}</math>

===Flotation===
Flotation is a process designed for specific solid-solid mixtures. It works by generating gas bubbles in a liquid that attach to selected solid particle. Afterwards, the particles rise to the liquid surface where they are removed by an overflow weir or mechanical scraper. The separation depends on the surface properties of the particles and its preference to attach to the gas bubbles. To meet the necessary requirements of the flotation process, a number of additives can be used to control things like the pH of the liquid-solid mixture, the activity of the solid surface, and the froth that can assist in separation. The bubbles can be produced by gaseous dispersion, dissolution, or electrolysis of the liquid (Peters & Timmerhaus, 2003).

===Centrifugation===
This process is similar to external field separation in that an external force field is applied to separate a mixture. When gravity separation is too slow due to particle densities, particle size, settling velocity, or the formation of an emulsion, centrifugation is commonly used. Centrifugal force increases the total force acting on the particle and results in faster separation times. This process is generally used to separate solids from liquids, however it can also be used to separate two liquids with very different densities (Peters & Timmerhaus, 2003).

===Drying===
Drying is performed to remove liquid from a liquid-solid mixture and produce a dry solid. Water is most often the liquid removed, but organic liquids are removed from solids on occasion as well. The heat required to vaporize the liquid is usually obtained by a series of gas-solid contacting devices. Feed condition and temperature sensitivity of the solid dictate the type of contacting device that is used. There are two groups of dryers that differ by the dependence of either mechanical means or fluid motion for gas solid contact. Another feature of dryers is to use either direct (hot gas) or indirect (conductive surface) heating (Peters & Timmerhaus, 2003).

===Evaporation===
Evaporators separate solvents from a solution by evaporation. The difference between evaporation and distillation is that evaporation requires the solute be nonvolatile. Because of this, a high separation can be achieved with one stage. Evaporators are essentially reboilers, so evaporation is a very energy-intensive process with a high thermal economy (Peters & Timmerhaus, 2003).

===Filtration===
Filtration is a process that separates a mixture of solid in a liquid or gas by passing the mixture through a porous medium in which the particles do not pass. Filtration is done by either cake filtration (particles found on the surface of the filter) or depth filtration (particles found within the filter). Cake filtration is generally performed with a cloth as the filtration medium (Peters & Timmerhaus, 2003).

==Conclusion==
Separation is a key part of most chemical processes, and there is a great variety of techniques to perform separation of compounds based on size, volatility, charge, and many other features. A common technique with which the process engineer should be familiar is distillation, but he or she should also be aware of the other available options. Some techniques may be less expensive, less energy-intensive, or more effective than distillation, depending on the specific separation problem. Therefore, the separation strategy should be carefully considered.

==References==
Air Control Technology Fact Sheet. Enviornmental Technology Fact Sheet. http://www3.epa.gov/ttncatc1/dir1/fcyclon.pdf

Belter PA, Cussler EL, Hu WS. Bioseparations: Downstream Processing for BIotechnology. New York: John Wiley; 1998.

Biegler LT, Grossmann IE, Westerberg AW. Systematic Methods of Chemical Process Design. Upper Saddle River: Prentice Hall; 1997.

Danckwerts P (1965) The Absorption of Gases in Liquids. Pure and Applied Chemistry UK 10:625-642.

Development Document for the Final Effluent Limitations Guidelines and Standards for the Metal Products and Machinery Point Source Category (Report). US Environmental Protection Agency. 2003.

Erwin, D. Industrial Chemical Process Design. New York: McGraw Hill, Professional Engineering; 2002.

Hage, David S. 1999Affinity Chromatography: A Review of Clinical Applications. Clinical Chemistry 45(5): 593–615.

Harrison RG, Todd P, Rudge SR, Petrides, DP. Bioseparations Science and Engineering. New York: Oxford University Press; 2003.

Hage, David S. 1999Affinity Chromatography: A Review of Clinical Applications. Clinical Chemistry 45(5): 593–615.

HisPur Cobalt Resin - Thermo Fisher Scientific N.d. https://www.thermofisher.com/order/catalog/product/89964, accessed February 21, 2016.

Lamella Plate Clarifier. Hydro International Web site. Available at: http://www.hydro-int.com/uk/products/lamella-plate-clarifier?s=0&r=uk. Accessed February 2, 2016.

Lean Oil Absorption. PetroGas Systems Web site. Available at: http://petrogassystems.com/technology/natural-gas-processing-and-dew-point-control/lean-oil-absorption. Accessed February 19, 2014.

Lichty, Jordan J., Joshua L. Malecki, Heather D. Agnew, Daniel J. Michelson-Horowitz, and Song Tan 2005Comparison of Affinity Tags for Protein Purification. Protein Expression and Purification 41(1): 98–105.

Merichem Gas Technologies. ®LO-CAT PROCESS available at http://www.merichem.com/images/casestudies/Desulfurization.pdf Accessed 6 Feb. 2015.

Miller L.N. & Zawacki T.S. , US 4080424, "Process for acid gas removal from gaseous mixtures", issued 21 Mar 1978, assigned to Institute of Gas Technology

NMED Surface Water Quality Bureau, New Mexico Water Systems Operator Certification Study Manual, New Mexico Environment Department, 2015.

Peters MS, Timmerhaus KD. Plant Design and Economics for Chemical Engineers. 5th ed. New York: McGraw Hill; 2003.

Prihasto, N; Lui, Q; Kim, S. Pre-treatment strategies for seawater desalination by reverse osmosis system. 2009; 249(1): 308-316. doi:10.1016/j.desal.2008.09.010

Schmidt Eberhard (2012) Waste Gases, Separation and Purification. Ullman’s Encyclopedia of Industrial Chemistry Germany 2:174-181.

Seider, W.D., Seader, J.D., and Lewin, D.R. (2004). ''Process Design Principles: Synthesis, Analysis, and Evaluation.'' New York: Wiley.

Stripping Column. Alfa Laval Web site. Available at: http://www.alfalaval.com/solution-finder/products/soft-column/Documents/Stripping%20Column.pdf. Accessed February 19, 2014.

Towler G, Sinnott R. Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013.

Turton, R.T., Bailie, R.C., Whiting, W.B., and Shaewitz, J.A. (2003). ''Analysis, Synthesis, and Design of Chemical Processes'' Upper Saddle River: Prentice-Hall.

Wankat, P.C. (2012). ''Separation Process Engineering.'' Upper Saddle River: Prentice-Hall.

Wilson, T.E., Clarifier Design, 2nd Ed., McGraw-Hill: New York, 2005.

Utility systems

2016-02-22T05:18:21Z

Bts511: /* Outcomes */

Authors: David Chen, [2014] Joshua Lee, [2015] Brett Sleyster, [2016] and Tom Aunins [2016] 

Stewards: David Chen, Jian Gong, and Fengqi You

Date Presented: January 13, 2014 /Date Revised: January 14, 2014

 


__TOC__

=Introduction=
Many chemical processes do not take place at ambient temperature or pressures. In order to reach these non-ambient conditions, utilities will have to be used to raise or lower temperatures and compress gases. (Towler, Towler/UOP) Utilities often contribute 5 to 10% of the price of a product, and may come from public or private utility companies or on-site plants. For purchased utilities, costs depend on consumption, while for company-owned utilities, there will be both capital and operating costs. They include things such as steam for heating, electricity, cooling water, refrigeration, fuels such as natural gas, wastewater treatment, waste disposal, and landfill. Steam is often the largest utility cost. Cogeneration unit can supply electricity accompanied with different steam pressures. (Seider 2009)

=Electricity=

Electricity is used to power many different kinds of equipment. It has many advantages: it is efficient (> 90%), reliable, available in a wide range of power, shaft speeds, designs, lifetimes, convenience, costs, and maintenance. It is generally used up to 200 hp, and sometimes over 10,000 Hp. In chemical process plants, the electricity demand is generally determined by the work or energy required for compression, pumping, air cooling, lights, and many other items. This electricity often times is purchased from local electricity providers, but many plants generate their own electricity via sophisticated processes.

Electricity is rarely used as a primary heat utility in large-scale chemical plants for a variety of reasons. The main disadvantages of using electricity as a heat utility are as follows (Towler 2012):

:*Heat from electricity is two to three times more expensive than heat from fuels. This is attributed to the lack of efficiency when creating heat from electricity.

:*Electrical heating units are expensive, require high maintenance, and must comply with strict safety regulations.

:*Electrical heating units are unsafe compared to steam heating units. In steam systems, the physically steam controls the temperature, whereas in electrical heating units temperature is controlled by temperature controllers, which can fail or burn out.

The use of electricity carries with it some hazards depending on the environment. Extra care must be taken when using electrically-powered equipment in areas which may have combustible fluids, vapors, or dust, and where liquids may be present. (Seider 2009 pg 606)

==Conventional Power Station==

In general, most electricity is generated from a conventional coal-fired process, whether it be on-site or purchased from a provider. Coal-fired processes have been used to create electricity throughout history, and technological advances have increased its efficiency and use worldwide. In a coal-fired steam station—much like a nuclear station—water is turned into steam, which in turn drives turbine generators to produce electricity. There are several variations on how to create energy from coal, but here are the basics of how a coal-fired process works:

[[File:coalfired.png|thumb|border|center|middle|upright=4|link=|atl=|General Coal-Fired Process Diagram (Duke Energy Company)]]

*Heat is created:
::Before the coal is burned, it is pulverized to the fineness of talcum powder. It is then mixed with hot air and blown into the firebox of the boiler. Burning in suspension, the coal/air mixture provides the most complete combustion and maximum heat possible.

*Water turns to steam:
::Highly purified water, pumped through pipes inside the boiler, is turned into steam by the heat. The steam reaches temperatures of up to 1,000 degrees Fahrenheit and pressures up to 3,500 pounds per square inch, and is piped to the turbine.

*Steam turns the turbine:
::The enormous pressure of the steam pushing against a series of giant turbine blades turns the turbine shaft. The turbine shaft is connected to the shaft of the generator, where magnets spin within wire coils to produce electricity.

*Steam is converted back to water:
::After doing its work in the turbine, the steam is drawn into a condenser, a large chamber in the basement of the power plant. In this important step, millions of gallons of cool water from a nearby source (such as a river or lake) are pumped through a network of tubes running through the condenser. The cool water in the tubes converts the steam back into water that can be used over and over again in the plant.

*Repeat:
::The cooling water is returned to its source without any contamination, and the steam water is returned to the boiler to repeat the cycle.

===Advantages & Disadvantages of Coal-Fired Energy Production===

[[File:coalfiredtable.png|thumb|border|center|middle|upright=4|link=|atl=|Advantages and Disadvantages of Coal-Fired Processes (Seider 2008)]]

==Gas-Turbine Cogeneration Process==

When generating energy on-site, many plants use a gas-turbine cogeneration process. The thermal efficiency of a gas-turbine process is in the range of 70-80% while conventional power stations, such as coal-fired processes, have a 30-40% efficiency. The lower efficiency in more conventional power stations is attributed to wasted heat in the exhaust steam in the condenser. One example of a gas-turbine process is outlined in the following figure. Figure 3.1 is a gas-turbine cogeneration process with a heat recovery steam generator (waste-heat) boiler.

[[File:cogeneration.png|thumb|border|center|middle|upright=4|link=|atl=|Gas-Turbine Cogeneration with a heat recovery steam generator boiler (Towler 2012 Fig 3.1)]]

Overall, the process illustrated is not much different from a coal-fired process. The main differences are that the cogeneration process creates both electricity and a heat utility, and cogeneration processes use natural gas instead of coal. Many of the advantages and disadvantages are similar to those of the coal-fired process, but the cogeneration has a much higher efficiency, creates heat to be used in another process, and uses a more volatile and expensive fuel. The main advantage of cogeneration over coal-fired energy production is that heat is not wasted. In coal-fired processes, heat is released and wasted during electricity generation. Cogeneration captures some, if not all of the byproduct for heat, which is an extremely useful utility that will be discussed in subsequent sections. In summary, the cogeneration plant is superior to the coal-fired process because of its higher efficiency and ability to create a useful heat utility.

Obviously any engineer would design the cogeneration plant to meet at least the energy requirement necessary for plant operation, but cogeneration plants often times are designed to exceed the plant electricity requirement to drive another source of capital. Many describe this scenario as a "make or buy" scenario (Towler 2012). This scenario provides chemical producers leverage when negotiating contracts with outsourced electricity providers and this allows plants to purchase electricity at a wholesale price. This is a huge advantage of considering on-site electricity production because electricity is needed in relatively high quantities for all chemical plants. Being able to minimize electricity costs, or even profit off of electricity production is a huge economical consideration that all plants employ.

=Process Heating=
The key objective of this section is to discuss how processes are heated. Heating utilities are necessary for proper usage of distillers, reactors, condensers, and several other integral types of equipment. More specifically, steam, fired heat, and hot oil/specialized heat transfer fluids will be discussed in the following subsections.

==Steam==
Steam is the most commonly used heat utility used in chemical plants, and as a result understanding how it is used is essential in the study of Utility systems. Steam is used both as a process fluid (feedstock, diluent to absorb heat of reaction, heating agent, and stripping agent in absorbers and adsorbers ) and utility. It can be used to drive pumps and compressors, ejectors (for producing a vacuum), and heat exchangers. As one can clearly see, steam is a versatile, and useful utility.

Here are a few advantages of using steam as opposed to other methods of process heating (Towler 2012):

:*By controlling the pressure of the steam, one can control the temperature at which the heat is released. Having a strong control over the temperature is essential in several processes. =
:*Steam is an efficient heat source because the heat of condensation of steam is very high. Meaning that there is is high output per mass of utility at a constant temperature.
:*Heat exchangers that use steam are relatively cheap because condensing steam has a high heat transfer coefficient.
:*Steam is nonflammable, nontoxic, and inert to several process fluids.

Chemical plants generally have a network of pipelines exclusively for providing steam. These networks generally have steam at a low pressure, a medium pressure, and a high pressure. The image below illustrates a basic steam system.

[[File:Steam.png|thumb|border|center|middle|upright=4|link=|atl=|(Towler 2012 Fig 3.2)]]

In the diagram above, boiler feed water at a high pressure is preheated and fed to other boilers. These other boilers superheat the steam to create a high pressure and high temperature steam stream. The steam is superheated past the dew point to account for heat loss in the pipelines. A portion of the high pressure steam is used for process heating in areas of the plant that require high temperatures. The rest of the high pressure steam is turned into medium pressure steam by valves and steam turbines. The medium pressure steam is then used to heat medium temperature processes and to form low pressure steam. The low pressure steam can be used to heat low pressure processes and it can be expanded in condensing turbines to create shaft work and energy. In summary, steam can be used for an innumerable amount of action items in a plant. High pressure, medium pressure, and low pressure steam can all be used as a heat source. Low pressure steam has utilities in creating electricity and it also has several other uses.

==Fired Heat==

In many cases, processes in a plant require a heat source stronger than high pressure and temperature steam. That is when fired heat is used, which is generally at temperatures above 523K. Streams can be heated directly in the furnace tubes or via a hot oil circuit or heat transfer fluid, which will be discussed in detail in the next section. Most fired heaters use natural gas as fuel because it burns cleaner than fuel oil. A cleaner burning fuel is always advantageous due to environmental and safety concerns. Furthermore, natural gases usually result in less wear and tear in burners and fuel lines.

Depending on the application of the fired-heater, different design specifications can be implemented to make the fired-heater as efficient as possible. The basic construction of a fired heater starts with a cylindrical chamber that is lined with with refractory bricks. The stream that is to be heated flows through tubes inside of the furnace. These tubes can be arranged in several different arrangements such as, around the walls of the furnace, or in horizontal or vertical banks (Towler 2012). The figure below illustrates the basic construction of the fired-heater and varying tube arrangements.

[[File:firedheater.png|thumb|border|center|middle|upright=4|link=|atl=|Basic Fired-Heater Diagrams (Towler 2012 Figure 19.66)]]

Fuel is burned to heat the entire furnace, and the heat transfer occurs from the combustion gases inside of the furnace across the tubes that are filled with our desired stream. The heat transfer between the tubes and the furnace is accomplished primarily via radiation. Modern designs take advantage of convective heat transfer by adding a smaller section on top of the fired-heater where the combustion gases flow over banks of tubes as seen in (c) in the figure above. Heat transfer can be further improved via convection by adding fins or pins in the combustion section.

The cost of fired heat can be calculated by the cost of fuel fired. Natural gas and heating oil are traded as commodities and prices can be found at many online trading sites or business news sites (i.e., www.cnn.money.com). Past historic prices for forecasting can be found in the Oil and Gas Journal or from the U.S. Energy Information Adminstration (www.eia.gov).

===Fuels===
Fuel is burned in utility facilities such as boilers, electricity generation, and cogeneration, and can be in solid, liquid, or gas form. It can also be burned to provide heating for a process or stream or to drive pumps and compressors. The fuel is usually burned with excess air to ensure complete combustion.

A way of quantifying the amount of heat generated is by using the heating values. Higher heating value (HHV) and the lower heating value (LHV) are used. The heating is the total heat evolved by complete combustion of a fuel with dry air with both at 60 ⁰F and the flue gas after combustion brought back down to ⁰F. If the water vapor in the flue gas is not condensed, we obtain the LHV. If the water vapor is condensed, the value of heat evolved is a bit higher, and this is the HHV.
Heating values for solids and liquids are usually on a per-mass basis, and gases on a per-volume basis.
A specified amount of heating cannot be met with the amount of fuel calculated using only the HHV. There will be heat losses, the flue gas temperature will be greater than 60 ⁰F, and water in the flue gas will typically be vapor. (Seider 608)

==Hot Oil/Specialized Heat Transfer Fluids==

Specialized heat transfer fluids and hot oil circuits are used as heat sources when steam and fired heat is not appropriate. Specialized heat transfer fluids and hot oil circuits are extremely versatile in that they can be used in the temperature range of 323K to 673K. This range however is quite variable. For hot oils, the upper temperature limit is gauged based off of the thermal decomposition of the oil and coking/fouling of heat exchanger tubes.

Hot oil circuit systems are most commonly used when the plant has several small temperature heating requirements because it is more economically sound. Rather than having several fired heaters heat each small temperature requirement, it is much more economical to have one fired heater heat the hot oil and circulate that oil through each of the process to meet all of the heat needs. Hot oil systems are also generally favored over high pressure steam in processes that involve high pressure differentials between the process stream and high pressure steam. Hot oil systems are favored in this scenario because of safety concerns. If the steam were to leak, the pressure drop could cause serious safety issues.

Mineral oils are the most commonly used heat transfer fluids, and one prominent example is Dowtherm A. Dowtherm A is a combination of 26.5 wt% diphenyl in diphenyl oxide (Towler 2012) and is extremely thermally stable. These mineral oil systems generally require high flow rates.

=Process Cooling=

==Cooling Water==
Cooling water is used to cool and/or condense streams. Cooling water is usually circulated between process heat exchangers and a cooling tower. Water is cooled during downward motion by contact with air blown upwards, which can bring the water temperature to come within ~ 5 ⁰F of air’s wet-bulb temperature.Approximately 80% of the temperature reduction is due to evaporation of the cooling water and heat transfer to the surrounding air. Water can also be cooled in spray ponds and cooling ponds. Both work by providing high area for water to exchange heat with air.
Water in cooling towers is lost through drift and blowdown, and makeup is usually 1.5 to 3% of the circulating rate.
If a large natural body of water is nearby, it can be used as a source of cooling water and discharged downstream. This water is usually filtered to remove salts and impurities that may lead to fouling, but it is not treated.

==Refrigeration==
Cooling water can usually be used to cool a stream to ~ 100 ⁰F. Air can only cool to ~ 120 ⁰F. Air may be used in places where water is scarce or more costly to transport. To cool or condense streams below 100 ⁰F, a refrigerant is typically used. Chilled brine can also be used, but is less common.

Until 1995, CFC Freon R-12 (dichlorodifuloromethane) and HCFC Freon R-22 (chlorodifuloromethane) were commonly used refrigerants. However, the chlorine atom in the molecules caused the depletion of the ozone layer. The U.S. Clean Air Act Amendments of 1990 went into effect in 1995, and the production of these refrigerants has since stopped or been greatly reduced.

Cost estimates are based on ton-day of refrigeration, where a ton is the heat that needs to be removed to freeze 1 ton per day of water at 32 ⁰F. Substitutes have since been developed. R-134a is often used in place of R-12. According to the EPA, R-134a is not combustible at ambient conditions, and is essentially non-toxic under 400 ppm, and is not ozone-depleting. (Seider pg 607)

=Energy Efficiency=
One of the chief concerns in selecting and designing process utility systems for heating and cooling is how to achieve the most energy efficient design. There are countless means by which plants lose energy, two of the foremost being through the mixing of different temperature or pressure streams and through the disposal of warmed cooling water. (Seider, 2009) Proper utilities design can help mitigate each of these losses as well as many others. The energy efficiency of a plant will depend primarily on the heating and cooling methods that are being used and the overall system design itself. These two parameters are important in determining how well energy is transferred to the desired media as well as how well that energy is recovered and recycled.

==Hot Utility Efficiency==
As mentioned above, the most commonly used utilities for process heating in large scale processes are steam, fired heat, and hot oil heaters. Of these, steam is the most commonly used. Electricity, while efficient at creating power, is not a viable source of heat in large industrial processes. Common ranges of heating efficiency for these three methods are displayed in Table 1. (Towler and Sinnott, 2012; Broughton, 1994)

{| class="wikitable"
|+Table 1: Process Heating Efficiencies
|-
! Process Heating Method
! Typical Efficiency
|-
| Steam (via coal boiler)
| 72%
|-
| Steam (via gas boiler)
| 73%
|-
| Steam (via oil boiler)
| 75%
|-
| Fired Heat w/ Convective Section
| 85%
|-
| Fired Heat w/o Convective Section
| 60%
|-
| Hot Oil Heaters/Vaporizers
| 80-85%
|}

===Steam===
As steam is so popular for heating purposes, it is useful to analyze the numerous ways in which losses can occur in steam systems. There are five primary sources of inefficiency and heat loss in the generation and distribution of steam throughout a process plant:

*The heat content of boiler exhaust gas
*Incomplete combustion of boiler fuel
*Radiant losses from the boiler exterior
*Blowdown losses
*Distribution losses (pipe transport, steam traps, etc.)

The first four of these losses take place at the boiler and contribute to the heating efficiencies seen in Table 1 for steam created with coal, gas, and oil. Several methods can be used to minimize these losses, one of the foremost being the control of air-to-fuel ratio in the boiler. This ratio must be managed by weighing losses due to uncombusted fuel against losses due to the heat content of excess exhaust gas. (Broughton, 1994)

[[File:Air-fuel ratio.PNG|thumb|border|center|middle|upright=4|link=|atl=|Air-to-fuel ratio relationship with heat losses (Broughton Fig 2.3)]]

The air-to-fuel ratio can be optimized using a feedback process controller. The control system will analyze the oxygen content of exhaust air and adjust the incoming air flow rate to achieve a set percentage of excess air. While desired excess oxygen will vary depending on the type of fuel, it is consistently seen that in the minimum loss-range a 1% increase in excess air will result in a 1% decrease in efficiency. ("Oxygen Control")

In addition to air-to-fuel ratio management, steam heat losses at the boiler can be mitigated via energy recovery, which is discussed in further detail [[Utility_systems#Energy_Recovery|below]].

Outside the boiler, losses in distribution of steam throughout facilities can decrease energy efficiency by a significant amount. This can cause up to a 60% increase in losses, but typically results in an overall steam efficiency of 50-55% (down from approximately 75% at the boiler). (Broughton, 1994) There are several ways that this issue can be addressed. First and foremost are steam metering systems, which can be used to monitor heat supply to separate sections of the process facility, check efficiency of fuel use, and determine which processes in a given plant have the highest steam requirements. Another method of minimizing transport losses is to decentralize steam generation systems. It can be advantageous to have numerous smaller boilers rather than a single centralized boiler so that steam does not have to travel as long of distances to reach its destination.

==Cold Utility Efficiency==
Efficiency in cooling processes depends more on the method used, and by extension the amount of coolant needed. Water and air utility efficiencies depend primarily on the fluid flow required to maintain the system at a desired temperature, while powered refrigeration utilities (for colder processes) have efficiencies at approximately 60%--but ranging up to 90%--of Carnot cycle efficiency, a metric of ideal refrigeration efficiency. (Towler and Sinnott, 2012) Cooling systems represent by definition a loss of energy from the main process to the utility stream, and as such it is often useful to find other uses for the heated media before discharge.

==Energy Recovery==
Recovery and recycle of energy is perhaps the most important aspect of creating an energy efficient plant design, and it is important for process engineers to fully consider possibilities for heat recovery in order to aid in economic viability.

===Process Heat Exchange===
Heat exchanger networks are a very common energy recovery method in industrial processes. These networks most frequently allow energy from heated product streams to be transferred to feed streams that must be brought up to process temperature. (Biegler, 1997) More information on the function and design of heat exchanger networks can be found on the [[Heat_exchanger#Heat_Exchanger_Networks|heat exchanger wiki page]]. The following are several examples of energy recovery via heat exchange that are used in industrial processes.

In distillation columns the bottoms and distillate effluents have the potential for energy exchange. Though the condenser at the top of the column cannot supply its waste heat to the reboiler due to their respective temperatures, the effluent streams can supply heat to the feed via a feed-effluent exchanger. This reduces the utility requirements to raise the feed to column temperature. (Biegler, 1997)

[[File:Feed_sterilization.jpg|thumb|border|right|middle|upright=4|link=|atl=|Feed sterilization schematic. (Towler and Sinnott Fig. 3.30)]]

Feed sterilization, commonly used in the food industry, is a common application for heat recovery through process stream heat exchange. In this application, the feed must be heated for a certain amount of time to kill any biological contaminants, after which it can be used to heat the new raw feed for sterilization. This reduces energy demands on the steam heater and thus reduces cost. (Towler and Sinnott, 2012)

In multi-vessel batch processes it can be advantageous to exchange heat as the process fluid is being transferred between vessels. Like the previous examples, this reduces the utility needed to bring the colder feed up to process temperature, thus reducing costs. (Towler and Sinnott, 2012)

[[File:Batch_heat_exchange.jpg|thumb|border|center|middle|upright=4|link=|atl=|Multi-vessel batch heat exchange schematic (Towler and Sinnott Fig. 3.31)]]

===Utility Regeneration===
[[File:Waste_heat_boiler.jpg|thumb|border|right|middle|200px|upright=4|link=|atl=|Industrial modular HRSG]]
When recovery of waste heat via transfer to other process streams is inconvenient or impossible, energy efficiency can still be improved through the regeneration of utilities. This is commonly done through the regeneration of steam by removing heat from exiting streams or from highly exothermic reactions. Waste heat in exiting streams can be removed via heat recovery steam generators (HRSGs), and is most often used on exiting gas streams. Heat recovery from reactions is a viable option when the reactor temperature will be at 150 C or above, as this will create steam at high enough pressure to be used in other processes. (Towler and Sinnott, 2012)

In the case of steam, waste heat and water treatment losses can be recovered from the utility generation process itself. One of the most common ways to do this is with an economizer. As seen in the schematic below, economizers have heat exchangers fit to the exhaust gas flow in order to transfer waste heat from these gases to the incoming boiler feed water. This can result in fuel energy savings of approximately 15% for typical excess air percentages in steam boilers. (Broughton, 1994; "Economizers")

[[File:Economizer.png|thumb|border|center|middle|upright=4|link=|atl=|Economizer schematic ("Economizers") and economizer energy recovery correlations (Broughton Fig 2.2)]]

[[File:Condensate recycle.PNG|thumb|border|right|middle|200px|upright=4|link=|atl=|Condensate return fuel savings (Broughton Fig 2.4)]]

Additionally, there are still more opportunities for recovery in steam generation processes. One typical case is the use of waste heat to preheat air entering the boiler. Optimistically, this can result in a 5% improvement in heat recovery for the system. Another method is the recycle of steam condensate back to the boiler. This has two main benefits: it retains the heat still present in the condensed steam and it retains the pretreatment chemicals added to the steam. (Broughton, 1994)

Furthermore, there is opportunity for energy recovery in the expansion of compressed gas through a turbine to create electricity, a process that can be economically viable given sufficiently high flows or pressure. One of the primary application of this type of energy recovery is in the creation of medium pressure and low pressure steam. In most processes, all steam is generated as high pressure steam and can be expanded through a turbine to decrease its pressure. Such technology has also been used in processes to synthesize ammonia, perform air separations, and synthesize nitric acid. (Towler and Sinnott, 2012) Recently, however, there has been a particularly strong interest for energy recovery in the natural gas industry, when gas is decompressed from major pipelines to residential low-pressure pipelines. A 2001 study estimated that there is the potential to recover 21 TWh, representing 11% of natural gas transport energy, via gas expansion. (Lehman)

=Process water and boiler-feed water=
Process water is water that will be directly used in the process. Boiler-feed water (BFW) is used to produce steam. Both may need to be purified to prevent impurities from contaminating a process or from foul equipment. It can be used as a cooling stream when the temperature of the stream to be cooled is greater than ~300 ⁰F. Cost of BFW can be partially offset by the steam credit.

Process water that undergoes moderate pretreatment can cost ~ $0.75/1,000 gal.

Extensive treatment ~ $6.00/1,000 gal.

Sterilized for pharmaceutical processes ~ $550/1,000 gal. (Seider pg 608)

=Demineralized Water=
In demineralized water, minerals have been removed by ion exchange. In boiler feed water, this is to prevent salt deposition, corrosion, formation of foam, and sluicing. In process water, the ions may contaminate the process.

=Waste Treatment=
Most chemical processes will produce some sort of waste. Disposal occurs to the atmosphere (in the case of some gases), sewers, body of water, or a landfill. Waste may require some treatment before disposal to meet regulations. Depending on process economics, byproducts may be recovered and processed. (Seider 2009 pg 609)

==Wastewater Treatment==
The United States EPA regulates industrial wastewater disposal through the Clean Water Act, introduced in 1948 as the Federal Water Pollution Control Act and amended to its current form in 1972. The sweeping 1972 amendments allowed the EPA to prevent industries and persons from discharging contaminated water into fresh water sources and set water quality standards. (Summary of the Clean Water Act) In accordance with this law, process plants in the United States treat wastewater at on-site or near-site treatment centers before releasing it into the surrounding environment.

Wastewater effluent streams, along with water runoff from around the plant, are treated to control for pH, toxicity, suspended solids, and biological oxygen demand (for aquatic life protection) prior to discharge. Each of these controls is typically addressed with a separate method. Acidity and basicity is balanced through the addition of an acid or alkaline solution. Toxic wastewater may be treated with chemical processes or simply diluted to safe concentrations. Suspended solids can be removed via filtration and/or with clarifiers. Oxygen demand of wastewater can be mitigated using activated sludge treatment processes. Once the water quality complies with the EPA, and state-mandated, regulations, it can be safely released. More information on the large number of industry-specific guidelines for waste effluent can be found on the EPA website (http://www.epa.gov/eg/industrial-effluent-guidelines).

=Air-Pollution Management=
===Introduction===

In the United States air pollution is regulated in the Clean Air Act, and almost all pollutant emitting plants are regulated under this law. The types of plants that can release significant emissions include petroleum refineries, sulfur recovery plants, carbon-black plants, fuel conversion plants, chemical process plants, fossil fuel plants, and petroleum storage and transfer facilities. To receive permission to construct a plant must undergo a review to show that it will not cause a violation of the Ambient Air Quality Standards(Peters, 2003).

==Methods==

There are two major types of pollutants that are released into the air, particulates and and gaseous pollutants. Particulates can be removed with mechanical forces while gaseous pollutants typically need to removed by chemical or physical means (Peters, 2003).

Particulates and volatile pollutants that need to be removed before disposal may be present. Particle removal equipment includes: cyclones, wet scrubbers, electrostatic precipitators, and fabric-filter systems (Seider 2009) such as bag filters (. The two charts below are from ''Plant Design and Economics for Chemical Engineers'' and show the types of equipment, separation methods, and particle sizes in different pollutant separation technologies (Peters, 1991).

[[File:AirPolChart.png|thumb|border|center|middle|upright=4|link=|atl=|Different Methods of Particle Separation, the Particle Sizes they Can Remove, and the Technologies Used]]

[[File:AirPolTable.png|thumb|border|center|middle|upright=4|link=|atl=| This Chart Shows the Specifications and Limitations of Different Separations Technologies Including Particle Size, Efficiencies, and Temperatures(Peters, 1991)]]

Methods for removing inorganic and organic gaseous pollutants include: absorption, adsorption, condensation, and combustion (Seider 2009 pg 609). A list of typical of gases pollutants and their sources from ''Plant Design and Economics for Chemical Engineers'' is shown below (Peters, 1991).

[[File:GaseousPollutants.png|thumb|border|center|middle|upright=4|link=|atl=| Common Gaseous Pollutants and their Sources]]

Typically Gas-liquid absorption processes are done completed in a vertical, countercurrent, flow through packed, plate, or spray towers. These systems require good liquid-gas contact and proper equipment. These systems also often have significant energy consumption because of large pressure drops (Peters, 2003). For high volume systems absorption by scrubbing with water or another solvent is the most widely used method (Towler, 2012). Dry adsorbents can be used to remove the last races of gaseous pollutants. Adsorption typically requires blowers, condensers, separators, and controls. You also typically need two packed beds so that one can be used while the other is regenerated. Examples of adsorbents are molecular sieves and activated carbon. Incineration is typically used when there are gas streams that have no recovery value. This can be done with direct flame or catalytic oxidation. Catalytic oxidation usually has higher capital costs, but lower operating costs because it does not require fuel.
===Wet Scrubbers===
Wet scrubber use lime or limestone and water to remove SO2 and acid gases. The mixture can be injected into a scrubber or the gases can be bubbled through this mixture. This results in removal of 90-98% of SO2 and and acid gases (Clean Coal Technologies).

===Dry Scrubbers===
Dry scrubber blow powdered adsorbents into a vessel with gases and then after it has captured the SO2 and acid gases it is separated from gas using a fabric filter. These systems remove 90-93% of the contaminants (Clean Coal Technologies).

===Low-NOx Burners===
The purpose of low NOx burners is to decrease the amount of NOx created when the coal is burned. This is done by injecting coal and air in boilers. This can result in 40-50% NOx reduction. If air is injected into the area above the burner this can actually cause almost 70% NOx reduction (Clean Coal Technologies).

===Selective Non-Catalytic and Catalytic Reductions===
These systems inject ammonia into gases to remove NOx. The Catalytic reductions add a catalyst to the ammonia being injected to aid in the removal. Con-catalytic reductions result in about 35% removal, but adding a catalyst can increase that amount to about 90%. The catalytic-reduction can remove up to 80% of mercury as well (Clean Coal Technologies).

===Fabric Filters===
Also known as baghouses,these [[Separation processes#Filtration| filters]] remove particulates by passing air through filters. These can separate as much as 99.9% of particulate matter (Clean Coal Technologies).

===Electrostatic Precipitators===
Electrostatic Precicpitators remove particulate matter as gas passes through a device that has charged metal plates. The particles are then removed because of static electricity. These systems can remove between 99-99.9% of particulate matter (Clean Coal Technologies).

===Super-critical Boilers===
Super-Critical Boilers and Ultra-Supercritical Boilers operate at temperature and pressure higher than regular boilers. By operating at higher temperature these systems become more efficient. Super Critical Boilers typically have 10%-20% CO2 emissions than other similar sub-critical technologies. Ultra-Supercritical boilers can be as much as 30% more efficent than sub-critical technologies (Clean Coal Technologies).

===More Information===

For more information regarding some of these separations equipment see [[Solids-involved equipment]]

For more information Regarding Cyclones see [[Separation processes#Cyclones| Cyclones]], and for modeling cyclones see their[[Solids-involved equipment#HYSYS Simulation |HYSYS Simulation]]

You can find more information on [[Separation processes#Absorption| Absorption]] and [[Separation processes#Adsorption| Adsorption]] in [[Separation processes]]

===Outcomes (Case Studies)===

====United States Common Pollutant Emissions====

The United States implemented the Clean Air Act in 1970 and since then emissions in the U.S. have been drastically reduced. As a matter of fact despite increases in the population in the last 35 years the amount of pollutants emitted have decreased by almost 70% (EPA). More importantly this demonstrates that reasonable efforts can be put towards environmental protection without causing too much harm to industry.

[[File:EPA.png|thumb|border|center|middle|upright=4|link=|atl=| GDP and Other Growth Factors Vs Common Pollutants Released]]

===China Clean Coal===

====Success Through 2005====

Coal is a very inexpensive and abundant source of energy and is abundant in China (Xu 2010). In China Coal the cause of 90% of SO2 emissions, 70% of dust emissions, and 67% of NOx emissions, and 70% of CO2 emissions. These numbers are staggering considering that Coal has never been more than 50 percent of China's Energy Supply (Xu 2010).

[[File:Coal.jpg|thumb|border|center|middle|upright=4|link=|atl=| Coal Consumption for Power and the Percentage of Coal Used (Xu 2010)]]

Despite the increasing coal consumption, high efficiency electric dust removal systems with efficiencies that are as high as 99.6% have greatly decreased soot emissions to 32% below 1980's levels as of 2005.

[[File:Soot.jpg|thumb|border|center|middle|upright=4|link=|atl=| SO2 and Soot Emissions from 1981 to 2005 (Xu 2010)]]

====Technologies and Life Cycle Comparison====
As of 2010 China was consuming 48.2% of coal globally. Four potential technologies integrated gasification combined cycle (IGCC), sub-critical coal power generation (Sub-C); super-critical coal power generation (Super-C) ultrasuper-critical coal power generation (USC). These technologies are compared mostly on the basis of net generating efficiency and efficiency. Net generating efficiency is the output of the plant divided by the total available energy in the fuel used (Liang, 2013).

where efficiency is define as: η=E/Etotal

and

Etotal=Emining+Etransportation+Egeneration

[[File:technologies.PNG|thumb|border|center|middle|upright=4|link=|atl=|A comparison of many of the parameters important to lfe-cycle analysis (Liang, 2013)]]

This demonstrates that USC has the best net generation efficiency, while also having the largest capacity for a single system.

[[File:Lifecycle2.PNG|thumb|border|center|middle|upright=4|link=|atl=| Efficiency Calculation (Liang, 2013)]]

This breaks down the energy efficiency of each system and demonstrates that USC is the most efficient

[[File:Lifecycle.PNG|thumb|border|center|middle|upright=4|link=|atl=| Various pollutant outputs (Liang, 2013)]]

IGCC and USC are lowest on different element, but since these all have different global warming potential it is difficult to tell which is the most efficient.

[[File:GWP.PNG|thumb|border|center|middle|upright=4|link=|atl=| Quantitative environemntal measures (Liang, 2013)]]

This table breaks down the different measures of which systems are better

[[File:Measures.PNG|thumb|border|center|middle|upright=4|link=|atl=| Environmental measures and their explanations (Liang, 2013)]]

And the table above defines all the relevant terms.

[[File:Price.PNG|thumb|border|center|middle|upright=4|link=|atl=| Capital cost in $/kW of clean coal power-generation technologies (Liang, 2013)]]

This demonstrates that the capital cost per unit of energy production is very high for IGCC, but shows that Super-C and USC are very competitive.

All of the data suggests that USC is the highest in both energy efficiency and net generating efficiency while competitive in price. The IGCC while potentially the best on emissions is too much more expensive to use on a wide scale (Liang, 2013).

=Solid Waste=
U.S. federal regulations require that solid waste be classified as hazardous or nonhazardous. Conditions for a classification of hazardous include: ignitability, corrosivity, reactivity, toxicity, or posing a substantial threat to the surrounding environment and its inhabitants. Hazardous waste must be treated on- or near-site before being removed in containers. Non-hazardous waste may be landfilled or incinerated in some cases. A typical estimate of costs for waste disposal is $0.03/lb for nonhazardous solids and $0.10/lb for hazardous solids. (Seider 2009 pg 609)

=References=

*Biegler LT, Grossmann IE, Westerberg AW. ''Systematic Methods of Chemical Process Design'' Prentice-Hall: Upper Saddle River, 1997.

*Broughton, J. ''Process Utility Systems: Introduction to Design, Operation and Maintenance'' Institution of Chemical Engineers: Rugby, Warwickshire, UK, 1994.

*Clean Coal Technologies. America's Coalition for Clean Coal Electricity. http://www.americaspower.org/clean-coal-technologies-1663/

*Duke Energy Company (2013). How Do Coal Fired Plants Work? Charlotte: Duke Energy.

*"Economizers." Energy Solutions Center Inc.: Boiler Consortium. http://www.cleanboiler.org/Eff_Improve/Efficiency/Economizers.asp. Accessed February 21, 2016.

*Lehman B, Worrell E. ''Electricity Production from Natural Gas Pressure Recovery Using Expansion Turbines.'' Lawrence Berkeley National Laboratory; 2001.

*"Overview of the Clean Air Act and Air Pollution." Environmental Protection Agency. November 17, 2015. http://www.epa.gov/clean-air-act-overview. Accessed February 5, 2016.

*"Oxygen Control." Energy Solutions Center Inc.: Boiler Consortium. http://www.cleanboiler.org/Eff_Improve/Efficiency/Oxygen_Control.asp. Accessed February 21, 2016.

*Peters, Max S.; Timmerhaus, Klaus D.; West, Ronald E. (2003). " Plant Design and Economics for Chemical Engineers." McGraw Hill Higher Education.

*Ruina Xu Energy (2010). "Clean coal technology development in China." Elsevier.

*Seider, Seader, Lewin, Widagdo. (2009). ''Plant Design and Economics for Chemical Engineers, 5th Edition.'' Hoboken: Wiley.

*Seider, Seader, Lewin. (2008). ''Product and Process Design Principles, 2nd Edition.'' Hoboken: Wiley.

*Summary of the Clean Water Act. United States EPA website. http://www.epa.gov/laws-regulations/summary-clean-water-act

*Towler, G.P. and Sinnot, R. (2012). ''Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design.'' Elsevier.

*Turton R.; Bailie, R.C.; Whiting, W.B.; Shaeiwitz J.A.; Bhattacharyya D. (2012). ''Analysis, Synthesis, and Design of Chemical Processes.'' Upper Saddle River: Prentice Hall.

*G.D. Ulrich, A Guide to Chemical Engineering Process Design and Economics, Wiley: New York, 1984.

*Xiaoye Liang, Zhihua Wang, Zhijun Zhou, , Zhenyu Huang, Junhu Zhou, and Kefa Cen. (2013). "Up-to-date life cycle assessment and comparison study of clean coal power generation technologies in China" Elsevier.

Utility systems

2016-02-22T05:17:48Z

Bts511: /* Technologies and Life Cycle Comparison */

Authors: David Chen, [2014] Joshua Lee, [2015] Brett Sleyster, [2016] and Tom Aunins [2016] 

Stewards: David Chen, Jian Gong, and Fengqi You

Date Presented: January 13, 2014 /Date Revised: January 14, 2014

 


__TOC__

=Introduction=
Many chemical processes do not take place at ambient temperature or pressures. In order to reach these non-ambient conditions, utilities will have to be used to raise or lower temperatures and compress gases. (Towler, Towler/UOP) Utilities often contribute 5 to 10% of the price of a product, and may come from public or private utility companies or on-site plants. For purchased utilities, costs depend on consumption, while for company-owned utilities, there will be both capital and operating costs. They include things such as steam for heating, electricity, cooling water, refrigeration, fuels such as natural gas, wastewater treatment, waste disposal, and landfill. Steam is often the largest utility cost. Cogeneration unit can supply electricity accompanied with different steam pressures. (Seider 2009)

=Electricity=

Electricity is used to power many different kinds of equipment. It has many advantages: it is efficient (> 90%), reliable, available in a wide range of power, shaft speeds, designs, lifetimes, convenience, costs, and maintenance. It is generally used up to 200 hp, and sometimes over 10,000 Hp. In chemical process plants, the electricity demand is generally determined by the work or energy required for compression, pumping, air cooling, lights, and many other items. This electricity often times is purchased from local electricity providers, but many plants generate their own electricity via sophisticated processes.

Electricity is rarely used as a primary heat utility in large-scale chemical plants for a variety of reasons. The main disadvantages of using electricity as a heat utility are as follows (Towler 2012):

:*Heat from electricity is two to three times more expensive than heat from fuels. This is attributed to the lack of efficiency when creating heat from electricity.

:*Electrical heating units are expensive, require high maintenance, and must comply with strict safety regulations.

:*Electrical heating units are unsafe compared to steam heating units. In steam systems, the physically steam controls the temperature, whereas in electrical heating units temperature is controlled by temperature controllers, which can fail or burn out.

The use of electricity carries with it some hazards depending on the environment. Extra care must be taken when using electrically-powered equipment in areas which may have combustible fluids, vapors, or dust, and where liquids may be present. (Seider 2009 pg 606)

==Conventional Power Station==

In general, most electricity is generated from a conventional coal-fired process, whether it be on-site or purchased from a provider. Coal-fired processes have been used to create electricity throughout history, and technological advances have increased its efficiency and use worldwide. In a coal-fired steam station—much like a nuclear station—water is turned into steam, which in turn drives turbine generators to produce electricity. There are several variations on how to create energy from coal, but here are the basics of how a coal-fired process works:

[[File:coalfired.png|thumb|border|center|middle|upright=4|link=|atl=|General Coal-Fired Process Diagram (Duke Energy Company)]]

*Heat is created:
::Before the coal is burned, it is pulverized to the fineness of talcum powder. It is then mixed with hot air and blown into the firebox of the boiler. Burning in suspension, the coal/air mixture provides the most complete combustion and maximum heat possible.

*Water turns to steam:
::Highly purified water, pumped through pipes inside the boiler, is turned into steam by the heat. The steam reaches temperatures of up to 1,000 degrees Fahrenheit and pressures up to 3,500 pounds per square inch, and is piped to the turbine.

*Steam turns the turbine:
::The enormous pressure of the steam pushing against a series of giant turbine blades turns the turbine shaft. The turbine shaft is connected to the shaft of the generator, where magnets spin within wire coils to produce electricity.

*Steam is converted back to water:
::After doing its work in the turbine, the steam is drawn into a condenser, a large chamber in the basement of the power plant. In this important step, millions of gallons of cool water from a nearby source (such as a river or lake) are pumped through a network of tubes running through the condenser. The cool water in the tubes converts the steam back into water that can be used over and over again in the plant.

*Repeat:
::The cooling water is returned to its source without any contamination, and the steam water is returned to the boiler to repeat the cycle.

===Advantages & Disadvantages of Coal-Fired Energy Production===

[[File:coalfiredtable.png|thumb|border|center|middle|upright=4|link=|atl=|Advantages and Disadvantages of Coal-Fired Processes (Seider 2008)]]

==Gas-Turbine Cogeneration Process==

When generating energy on-site, many plants use a gas-turbine cogeneration process. The thermal efficiency of a gas-turbine process is in the range of 70-80% while conventional power stations, such as coal-fired processes, have a 30-40% efficiency. The lower efficiency in more conventional power stations is attributed to wasted heat in the exhaust steam in the condenser. One example of a gas-turbine process is outlined in the following figure. Figure 3.1 is a gas-turbine cogeneration process with a heat recovery steam generator (waste-heat) boiler.

[[File:cogeneration.png|thumb|border|center|middle|upright=4|link=|atl=|Gas-Turbine Cogeneration with a heat recovery steam generator boiler (Towler 2012 Fig 3.1)]]

Overall, the process illustrated is not much different from a coal-fired process. The main differences are that the cogeneration process creates both electricity and a heat utility, and cogeneration processes use natural gas instead of coal. Many of the advantages and disadvantages are similar to those of the coal-fired process, but the cogeneration has a much higher efficiency, creates heat to be used in another process, and uses a more volatile and expensive fuel. The main advantage of cogeneration over coal-fired energy production is that heat is not wasted. In coal-fired processes, heat is released and wasted during electricity generation. Cogeneration captures some, if not all of the byproduct for heat, which is an extremely useful utility that will be discussed in subsequent sections. In summary, the cogeneration plant is superior to the coal-fired process because of its higher efficiency and ability to create a useful heat utility.

Obviously any engineer would design the cogeneration plant to meet at least the energy requirement necessary for plant operation, but cogeneration plants often times are designed to exceed the plant electricity requirement to drive another source of capital. Many describe this scenario as a "make or buy" scenario (Towler 2012). This scenario provides chemical producers leverage when negotiating contracts with outsourced electricity providers and this allows plants to purchase electricity at a wholesale price. This is a huge advantage of considering on-site electricity production because electricity is needed in relatively high quantities for all chemical plants. Being able to minimize electricity costs, or even profit off of electricity production is a huge economical consideration that all plants employ.

=Process Heating=
The key objective of this section is to discuss how processes are heated. Heating utilities are necessary for proper usage of distillers, reactors, condensers, and several other integral types of equipment. More specifically, steam, fired heat, and hot oil/specialized heat transfer fluids will be discussed in the following subsections.

==Steam==
Steam is the most commonly used heat utility used in chemical plants, and as a result understanding how it is used is essential in the study of Utility systems. Steam is used both as a process fluid (feedstock, diluent to absorb heat of reaction, heating agent, and stripping agent in absorbers and adsorbers ) and utility. It can be used to drive pumps and compressors, ejectors (for producing a vacuum), and heat exchangers. As one can clearly see, steam is a versatile, and useful utility.

Here are a few advantages of using steam as opposed to other methods of process heating (Towler 2012):

:*By controlling the pressure of the steam, one can control the temperature at which the heat is released. Having a strong control over the temperature is essential in several processes. =
:*Steam is an efficient heat source because the heat of condensation of steam is very high. Meaning that there is is high output per mass of utility at a constant temperature.
:*Heat exchangers that use steam are relatively cheap because condensing steam has a high heat transfer coefficient.
:*Steam is nonflammable, nontoxic, and inert to several process fluids.

Chemical plants generally have a network of pipelines exclusively for providing steam. These networks generally have steam at a low pressure, a medium pressure, and a high pressure. The image below illustrates a basic steam system.

[[File:Steam.png|thumb|border|center|middle|upright=4|link=|atl=|(Towler 2012 Fig 3.2)]]

In the diagram above, boiler feed water at a high pressure is preheated and fed to other boilers. These other boilers superheat the steam to create a high pressure and high temperature steam stream. The steam is superheated past the dew point to account for heat loss in the pipelines. A portion of the high pressure steam is used for process heating in areas of the plant that require high temperatures. The rest of the high pressure steam is turned into medium pressure steam by valves and steam turbines. The medium pressure steam is then used to heat medium temperature processes and to form low pressure steam. The low pressure steam can be used to heat low pressure processes and it can be expanded in condensing turbines to create shaft work and energy. In summary, steam can be used for an innumerable amount of action items in a plant. High pressure, medium pressure, and low pressure steam can all be used as a heat source. Low pressure steam has utilities in creating electricity and it also has several other uses.

==Fired Heat==

In many cases, processes in a plant require a heat source stronger than high pressure and temperature steam. That is when fired heat is used, which is generally at temperatures above 523K. Streams can be heated directly in the furnace tubes or via a hot oil circuit or heat transfer fluid, which will be discussed in detail in the next section. Most fired heaters use natural gas as fuel because it burns cleaner than fuel oil. A cleaner burning fuel is always advantageous due to environmental and safety concerns. Furthermore, natural gases usually result in less wear and tear in burners and fuel lines.

Depending on the application of the fired-heater, different design specifications can be implemented to make the fired-heater as efficient as possible. The basic construction of a fired heater starts with a cylindrical chamber that is lined with with refractory bricks. The stream that is to be heated flows through tubes inside of the furnace. These tubes can be arranged in several different arrangements such as, around the walls of the furnace, or in horizontal or vertical banks (Towler 2012). The figure below illustrates the basic construction of the fired-heater and varying tube arrangements.

[[File:firedheater.png|thumb|border|center|middle|upright=4|link=|atl=|Basic Fired-Heater Diagrams (Towler 2012 Figure 19.66)]]

Fuel is burned to heat the entire furnace, and the heat transfer occurs from the combustion gases inside of the furnace across the tubes that are filled with our desired stream. The heat transfer between the tubes and the furnace is accomplished primarily via radiation. Modern designs take advantage of convective heat transfer by adding a smaller section on top of the fired-heater where the combustion gases flow over banks of tubes as seen in (c) in the figure above. Heat transfer can be further improved via convection by adding fins or pins in the combustion section.

The cost of fired heat can be calculated by the cost of fuel fired. Natural gas and heating oil are traded as commodities and prices can be found at many online trading sites or business news sites (i.e., www.cnn.money.com). Past historic prices for forecasting can be found in the Oil and Gas Journal or from the U.S. Energy Information Adminstration (www.eia.gov).

===Fuels===
Fuel is burned in utility facilities such as boilers, electricity generation, and cogeneration, and can be in solid, liquid, or gas form. It can also be burned to provide heating for a process or stream or to drive pumps and compressors. The fuel is usually burned with excess air to ensure complete combustion.

A way of quantifying the amount of heat generated is by using the heating values. Higher heating value (HHV) and the lower heating value (LHV) are used. The heating is the total heat evolved by complete combustion of a fuel with dry air with both at 60 ⁰F and the flue gas after combustion brought back down to ⁰F. If the water vapor in the flue gas is not condensed, we obtain the LHV. If the water vapor is condensed, the value of heat evolved is a bit higher, and this is the HHV.
Heating values for solids and liquids are usually on a per-mass basis, and gases on a per-volume basis.
A specified amount of heating cannot be met with the amount of fuel calculated using only the HHV. There will be heat losses, the flue gas temperature will be greater than 60 ⁰F, and water in the flue gas will typically be vapor. (Seider 608)

==Hot Oil/Specialized Heat Transfer Fluids==

Specialized heat transfer fluids and hot oil circuits are used as heat sources when steam and fired heat is not appropriate. Specialized heat transfer fluids and hot oil circuits are extremely versatile in that they can be used in the temperature range of 323K to 673K. This range however is quite variable. For hot oils, the upper temperature limit is gauged based off of the thermal decomposition of the oil and coking/fouling of heat exchanger tubes.

Hot oil circuit systems are most commonly used when the plant has several small temperature heating requirements because it is more economically sound. Rather than having several fired heaters heat each small temperature requirement, it is much more economical to have one fired heater heat the hot oil and circulate that oil through each of the process to meet all of the heat needs. Hot oil systems are also generally favored over high pressure steam in processes that involve high pressure differentials between the process stream and high pressure steam. Hot oil systems are favored in this scenario because of safety concerns. If the steam were to leak, the pressure drop could cause serious safety issues.

Mineral oils are the most commonly used heat transfer fluids, and one prominent example is Dowtherm A. Dowtherm A is a combination of 26.5 wt% diphenyl in diphenyl oxide (Towler 2012) and is extremely thermally stable. These mineral oil systems generally require high flow rates.

=Process Cooling=

==Cooling Water==
Cooling water is used to cool and/or condense streams. Cooling water is usually circulated between process heat exchangers and a cooling tower. Water is cooled during downward motion by contact with air blown upwards, which can bring the water temperature to come within ~ 5 ⁰F of air’s wet-bulb temperature.Approximately 80% of the temperature reduction is due to evaporation of the cooling water and heat transfer to the surrounding air. Water can also be cooled in spray ponds and cooling ponds. Both work by providing high area for water to exchange heat with air.
Water in cooling towers is lost through drift and blowdown, and makeup is usually 1.5 to 3% of the circulating rate.
If a large natural body of water is nearby, it can be used as a source of cooling water and discharged downstream. This water is usually filtered to remove salts and impurities that may lead to fouling, but it is not treated.

==Refrigeration==
Cooling water can usually be used to cool a stream to ~ 100 ⁰F. Air can only cool to ~ 120 ⁰F. Air may be used in places where water is scarce or more costly to transport. To cool or condense streams below 100 ⁰F, a refrigerant is typically used. Chilled brine can also be used, but is less common.

Until 1995, CFC Freon R-12 (dichlorodifuloromethane) and HCFC Freon R-22 (chlorodifuloromethane) were commonly used refrigerants. However, the chlorine atom in the molecules caused the depletion of the ozone layer. The U.S. Clean Air Act Amendments of 1990 went into effect in 1995, and the production of these refrigerants has since stopped or been greatly reduced.

Cost estimates are based on ton-day of refrigeration, where a ton is the heat that needs to be removed to freeze 1 ton per day of water at 32 ⁰F. Substitutes have since been developed. R-134a is often used in place of R-12. According to the EPA, R-134a is not combustible at ambient conditions, and is essentially non-toxic under 400 ppm, and is not ozone-depleting. (Seider pg 607)

=Energy Efficiency=
One of the chief concerns in selecting and designing process utility systems for heating and cooling is how to achieve the most energy efficient design. There are countless means by which plants lose energy, two of the foremost being through the mixing of different temperature or pressure streams and through the disposal of warmed cooling water. (Seider, 2009) Proper utilities design can help mitigate each of these losses as well as many others. The energy efficiency of a plant will depend primarily on the heating and cooling methods that are being used and the overall system design itself. These two parameters are important in determining how well energy is transferred to the desired media as well as how well that energy is recovered and recycled.

==Hot Utility Efficiency==
As mentioned above, the most commonly used utilities for process heating in large scale processes are steam, fired heat, and hot oil heaters. Of these, steam is the most commonly used. Electricity, while efficient at creating power, is not a viable source of heat in large industrial processes. Common ranges of heating efficiency for these three methods are displayed in Table 1. (Towler and Sinnott, 2012; Broughton, 1994)

{| class="wikitable"
|+Table 1: Process Heating Efficiencies
|-
! Process Heating Method
! Typical Efficiency
|-
| Steam (via coal boiler)
| 72%
|-
| Steam (via gas boiler)
| 73%
|-
| Steam (via oil boiler)
| 75%
|-
| Fired Heat w/ Convective Section
| 85%
|-
| Fired Heat w/o Convective Section
| 60%
|-
| Hot Oil Heaters/Vaporizers
| 80-85%
|}

===Steam===
As steam is so popular for heating purposes, it is useful to analyze the numerous ways in which losses can occur in steam systems. There are five primary sources of inefficiency and heat loss in the generation and distribution of steam throughout a process plant:

*The heat content of boiler exhaust gas
*Incomplete combustion of boiler fuel
*Radiant losses from the boiler exterior
*Blowdown losses
*Distribution losses (pipe transport, steam traps, etc.)

The first four of these losses take place at the boiler and contribute to the heating efficiencies seen in Table 1 for steam created with coal, gas, and oil. Several methods can be used to minimize these losses, one of the foremost being the control of air-to-fuel ratio in the boiler. This ratio must be managed by weighing losses due to uncombusted fuel against losses due to the heat content of excess exhaust gas. (Broughton, 1994)

[[File:Air-fuel ratio.PNG|thumb|border|center|middle|upright=4|link=|atl=|Air-to-fuel ratio relationship with heat losses (Broughton Fig 2.3)]]

The air-to-fuel ratio can be optimized using a feedback process controller. The control system will analyze the oxygen content of exhaust air and adjust the incoming air flow rate to achieve a set percentage of excess air. While desired excess oxygen will vary depending on the type of fuel, it is consistently seen that in the minimum loss-range a 1% increase in excess air will result in a 1% decrease in efficiency. ("Oxygen Control")

In addition to air-to-fuel ratio management, steam heat losses at the boiler can be mitigated via energy recovery, which is discussed in further detail [[Utility_systems#Energy_Recovery|below]].

Outside the boiler, losses in distribution of steam throughout facilities can decrease energy efficiency by a significant amount. This can cause up to a 60% increase in losses, but typically results in an overall steam efficiency of 50-55% (down from approximately 75% at the boiler). (Broughton, 1994) There are several ways that this issue can be addressed. First and foremost are steam metering systems, which can be used to monitor heat supply to separate sections of the process facility, check efficiency of fuel use, and determine which processes in a given plant have the highest steam requirements. Another method of minimizing transport losses is to decentralize steam generation systems. It can be advantageous to have numerous smaller boilers rather than a single centralized boiler so that steam does not have to travel as long of distances to reach its destination.

==Cold Utility Efficiency==
Efficiency in cooling processes depends more on the method used, and by extension the amount of coolant needed. Water and air utility efficiencies depend primarily on the fluid flow required to maintain the system at a desired temperature, while powered refrigeration utilities (for colder processes) have efficiencies at approximately 60%--but ranging up to 90%--of Carnot cycle efficiency, a metric of ideal refrigeration efficiency. (Towler and Sinnott, 2012) Cooling systems represent by definition a loss of energy from the main process to the utility stream, and as such it is often useful to find other uses for the heated media before discharge.

==Energy Recovery==
Recovery and recycle of energy is perhaps the most important aspect of creating an energy efficient plant design, and it is important for process engineers to fully consider possibilities for heat recovery in order to aid in economic viability.

===Process Heat Exchange===
Heat exchanger networks are a very common energy recovery method in industrial processes. These networks most frequently allow energy from heated product streams to be transferred to feed streams that must be brought up to process temperature. (Biegler, 1997) More information on the function and design of heat exchanger networks can be found on the [[Heat_exchanger#Heat_Exchanger_Networks|heat exchanger wiki page]]. The following are several examples of energy recovery via heat exchange that are used in industrial processes.

In distillation columns the bottoms and distillate effluents have the potential for energy exchange. Though the condenser at the top of the column cannot supply its waste heat to the reboiler due to their respective temperatures, the effluent streams can supply heat to the feed via a feed-effluent exchanger. This reduces the utility requirements to raise the feed to column temperature. (Biegler, 1997)

[[File:Feed_sterilization.jpg|thumb|border|right|middle|upright=4|link=|atl=|Feed sterilization schematic. (Towler and Sinnott Fig. 3.30)]]

Feed sterilization, commonly used in the food industry, is a common application for heat recovery through process stream heat exchange. In this application, the feed must be heated for a certain amount of time to kill any biological contaminants, after which it can be used to heat the new raw feed for sterilization. This reduces energy demands on the steam heater and thus reduces cost. (Towler and Sinnott, 2012)

In multi-vessel batch processes it can be advantageous to exchange heat as the process fluid is being transferred between vessels. Like the previous examples, this reduces the utility needed to bring the colder feed up to process temperature, thus reducing costs. (Towler and Sinnott, 2012)

[[File:Batch_heat_exchange.jpg|thumb|border|center|middle|upright=4|link=|atl=|Multi-vessel batch heat exchange schematic (Towler and Sinnott Fig. 3.31)]]

===Utility Regeneration===
[[File:Waste_heat_boiler.jpg|thumb|border|right|middle|200px|upright=4|link=|atl=|Industrial modular HRSG]]
When recovery of waste heat via transfer to other process streams is inconvenient or impossible, energy efficiency can still be improved through the regeneration of utilities. This is commonly done through the regeneration of steam by removing heat from exiting streams or from highly exothermic reactions. Waste heat in exiting streams can be removed via heat recovery steam generators (HRSGs), and is most often used on exiting gas streams. Heat recovery from reactions is a viable option when the reactor temperature will be at 150 C or above, as this will create steam at high enough pressure to be used in other processes. (Towler and Sinnott, 2012)

In the case of steam, waste heat and water treatment losses can be recovered from the utility generation process itself. One of the most common ways to do this is with an economizer. As seen in the schematic below, economizers have heat exchangers fit to the exhaust gas flow in order to transfer waste heat from these gases to the incoming boiler feed water. This can result in fuel energy savings of approximately 15% for typical excess air percentages in steam boilers. (Broughton, 1994; "Economizers")

[[File:Economizer.png|thumb|border|center|middle|upright=4|link=|atl=|Economizer schematic ("Economizers") and economizer energy recovery correlations (Broughton Fig 2.2)]]

[[File:Condensate recycle.PNG|thumb|border|right|middle|200px|upright=4|link=|atl=|Condensate return fuel savings (Broughton Fig 2.4)]]

Additionally, there are still more opportunities for recovery in steam generation processes. One typical case is the use of waste heat to preheat air entering the boiler. Optimistically, this can result in a 5% improvement in heat recovery for the system. Another method is the recycle of steam condensate back to the boiler. This has two main benefits: it retains the heat still present in the condensed steam and it retains the pretreatment chemicals added to the steam. (Broughton, 1994)

Furthermore, there is opportunity for energy recovery in the expansion of compressed gas through a turbine to create electricity, a process that can be economically viable given sufficiently high flows or pressure. One of the primary application of this type of energy recovery is in the creation of medium pressure and low pressure steam. In most processes, all steam is generated as high pressure steam and can be expanded through a turbine to decrease its pressure. Such technology has also been used in processes to synthesize ammonia, perform air separations, and synthesize nitric acid. (Towler and Sinnott, 2012) Recently, however, there has been a particularly strong interest for energy recovery in the natural gas industry, when gas is decompressed from major pipelines to residential low-pressure pipelines. A 2001 study estimated that there is the potential to recover 21 TWh, representing 11% of natural gas transport energy, via gas expansion. (Lehman)

=Process water and boiler-feed water=
Process water is water that will be directly used in the process. Boiler-feed water (BFW) is used to produce steam. Both may need to be purified to prevent impurities from contaminating a process or from foul equipment. It can be used as a cooling stream when the temperature of the stream to be cooled is greater than ~300 ⁰F. Cost of BFW can be partially offset by the steam credit.

Process water that undergoes moderate pretreatment can cost ~ $0.75/1,000 gal.

Extensive treatment ~ $6.00/1,000 gal.

Sterilized for pharmaceutical processes ~ $550/1,000 gal. (Seider pg 608)

=Demineralized Water=
In demineralized water, minerals have been removed by ion exchange. In boiler feed water, this is to prevent salt deposition, corrosion, formation of foam, and sluicing. In process water, the ions may contaminate the process.

=Waste Treatment=
Most chemical processes will produce some sort of waste. Disposal occurs to the atmosphere (in the case of some gases), sewers, body of water, or a landfill. Waste may require some treatment before disposal to meet regulations. Depending on process economics, byproducts may be recovered and processed. (Seider 2009 pg 609)

==Wastewater Treatment==
The United States EPA regulates industrial wastewater disposal through the Clean Water Act, introduced in 1948 as the Federal Water Pollution Control Act and amended to its current form in 1972. The sweeping 1972 amendments allowed the EPA to prevent industries and persons from discharging contaminated water into fresh water sources and set water quality standards. (Summary of the Clean Water Act) In accordance with this law, process plants in the United States treat wastewater at on-site or near-site treatment centers before releasing it into the surrounding environment.

Wastewater effluent streams, along with water runoff from around the plant, are treated to control for pH, toxicity, suspended solids, and biological oxygen demand (for aquatic life protection) prior to discharge. Each of these controls is typically addressed with a separate method. Acidity and basicity is balanced through the addition of an acid or alkaline solution. Toxic wastewater may be treated with chemical processes or simply diluted to safe concentrations. Suspended solids can be removed via filtration and/or with clarifiers. Oxygen demand of wastewater can be mitigated using activated sludge treatment processes. Once the water quality complies with the EPA, and state-mandated, regulations, it can be safely released. More information on the large number of industry-specific guidelines for waste effluent can be found on the EPA website (http://www.epa.gov/eg/industrial-effluent-guidelines).

=Air-Pollution Management=
===Introduction===

In the United States air pollution is regulated in the Clean Air Act, and almost all pollutant emitting plants are regulated under this law. The types of plants that can release significant emissions include petroleum refineries, sulfur recovery plants, carbon-black plants, fuel conversion plants, chemical process plants, fossil fuel plants, and petroleum storage and transfer facilities. To receive permission to construct a plant must undergo a review to show that it will not cause a violation of the Ambient Air Quality Standards(Peters, 2003).

==Methods==

There are two major types of pollutants that are released into the air, particulates and and gaseous pollutants. Particulates can be removed with mechanical forces while gaseous pollutants typically need to removed by chemical or physical means (Peters, 2003).

Particulates and volatile pollutants that need to be removed before disposal may be present. Particle removal equipment includes: cyclones, wet scrubbers, electrostatic precipitators, and fabric-filter systems (Seider 2009) such as bag filters (. The two charts below are from ''Plant Design and Economics for Chemical Engineers'' and show the types of equipment, separation methods, and particle sizes in different pollutant separation technologies (Peters, 1991).

[[File:AirPolChart.png|thumb|border|center|middle|upright=4|link=|atl=|Different Methods of Particle Separation, the Particle Sizes they Can Remove, and the Technologies Used]]

[[File:AirPolTable.png|thumb|border|center|middle|upright=4|link=|atl=| This Chart Shows the Specifications and Limitations of Different Separations Technologies Including Particle Size, Efficiencies, and Temperatures(Peters, 1991)]]

Methods for removing inorganic and organic gaseous pollutants include: absorption, adsorption, condensation, and combustion (Seider 2009 pg 609). A list of typical of gases pollutants and their sources from ''Plant Design and Economics for Chemical Engineers'' is shown below (Peters, 1991).

[[File:GaseousPollutants.png|thumb|border|center|middle|upright=4|link=|atl=| Common Gaseous Pollutants and their Sources]]

Typically Gas-liquid absorption processes are done completed in a vertical, countercurrent, flow through packed, plate, or spray towers. These systems require good liquid-gas contact and proper equipment. These systems also often have significant energy consumption because of large pressure drops (Peters, 2003). For high volume systems absorption by scrubbing with water or another solvent is the most widely used method (Towler, 2012). Dry adsorbents can be used to remove the last races of gaseous pollutants. Adsorption typically requires blowers, condensers, separators, and controls. You also typically need two packed beds so that one can be used while the other is regenerated. Examples of adsorbents are molecular sieves and activated carbon. Incineration is typically used when there are gas streams that have no recovery value. This can be done with direct flame or catalytic oxidation. Catalytic oxidation usually has higher capital costs, but lower operating costs because it does not require fuel.
===Wet Scrubbers===
Wet scrubber use lime or limestone and water to remove SO2 and acid gases. The mixture can be injected into a scrubber or the gases can be bubbled through this mixture. This results in removal of 90-98% of SO2 and and acid gases (Clean Coal Technologies).

===Dry Scrubbers===
Dry scrubber blow powdered adsorbents into a vessel with gases and then after it has captured the SO2 and acid gases it is separated from gas using a fabric filter. These systems remove 90-93% of the contaminants (Clean Coal Technologies).

===Low-NOx Burners===
The purpose of low NOx burners is to decrease the amount of NOx created when the coal is burned. This is done by injecting coal and air in boilers. This can result in 40-50% NOx reduction. If air is injected into the area above the burner this can actually cause almost 70% NOx reduction (Clean Coal Technologies).

===Selective Non-Catalytic and Catalytic Reductions===
These systems inject ammonia into gases to remove NOx. The Catalytic reductions add a catalyst to the ammonia being injected to aid in the removal. Con-catalytic reductions result in about 35% removal, but adding a catalyst can increase that amount to about 90%. The catalytic-reduction can remove up to 80% of mercury as well (Clean Coal Technologies).

===Fabric Filters===
Also known as baghouses,these [[Separation processes#Filtration| filters]] remove particulates by passing air through filters. These can separate as much as 99.9% of particulate matter (Clean Coal Technologies).

===Electrostatic Precipitators===
Electrostatic Precicpitators remove particulate matter as gas passes through a device that has charged metal plates. The particles are then removed because of static electricity. These systems can remove between 99-99.9% of particulate matter (Clean Coal Technologies).

===Super-critical Boilers===
Super-Critical Boilers and Ultra-Supercritical Boilers operate at temperature and pressure higher than regular boilers. By operating at higher temperature these systems become more efficient. Super Critical Boilers typically have 10%-20% CO2 emissions than other similar sub-critical technologies. Ultra-Supercritical boilers can be as much as 30% more efficent than sub-critical technologies (Clean Coal Technologies).

===More Information===

For more information regarding some of these separations equipment see [[Solids-involved equipment]]

For more information Regarding Cyclones see [[Separation processes#Cyclones| Cyclones]], and for modeling cyclones see their[[Solids-involved equipment#HYSYS Simulation |HYSYS Simulation]]

You can find more information on [[Separation processes#Absorption| Absorption]] and [[Separation processes#Adsorption| Adsorption]] in [[Separation processes]]

===Outcomes===

====United States Common Pollutant Emissions====

The United States implemented the Clean Air Act in 1970 and since then emissions in the U.S. have been drastically reduced. As a matter of fact despite increases in the population in the last 35 years the amount of pollutants emitted have decreased by almost 70% (EPA). More importantly this demonstrates that reasonable efforts can be put towards environmental protection without causing too much harm to industry.

[[File:EPA.png|thumb|border|center|middle|upright=4|link=|atl=| GDP and Other Growth Factors Vs Common Pollutants Released]]

===China Clean Coal===

====Success Through 2005====

Coal is a very inexpensive and abundant source of energy and is abundant in China (Xu 2010). In China Coal the cause of 90% of SO2 emissions, 70% of dust emissions, and 67% of NOx emissions, and 70% of CO2 emissions. These numbers are staggering considering that Coal has never been more than 50 percent of China's Energy Supply (Xu 2010).

[[File:Coal.jpg|thumb|border|center|middle|upright=4|link=|atl=| Coal Consumption for Power and the Percentage of Coal Used (Xu 2010)]]

Despite the increasing coal consumption, high efficiency electric dust removal systems with efficiencies that are as high as 99.6% have greatly decreased soot emissions to 32% below 1980's levels as of 2005.

[[File:Soot.jpg|thumb|border|center|middle|upright=4|link=|atl=| SO2 and Soot Emissions from 1981 to 2005 (Xu 2010)]]

====Technologies and Life Cycle Comparison====
As of 2010 China was consuming 48.2% of coal globally. Four potential technologies integrated gasification combined cycle (IGCC), sub-critical coal power generation (Sub-C); super-critical coal power generation (Super-C) ultrasuper-critical coal power generation (USC). These technologies are compared mostly on the basis of net generating efficiency and efficiency. Net generating efficiency is the output of the plant divided by the total available energy in the fuel used (Liang, 2013).

where efficiency is define as: η=E/Etotal

and

Etotal=Emining+Etransportation+Egeneration

[[File:technologies.PNG|thumb|border|center|middle|upright=4|link=|atl=|A comparison of many of the parameters important to lfe-cycle analysis (Liang, 2013)]]

This demonstrates that USC has the best net generation efficiency, while also having the largest capacity for a single system.

[[File:Lifecycle2.PNG|thumb|border|center|middle|upright=4|link=|atl=| Efficiency Calculation (Liang, 2013)]]

This breaks down the energy efficiency of each system and demonstrates that USC is the most efficient

[[File:Lifecycle.PNG|thumb|border|center|middle|upright=4|link=|atl=| Various pollutant outputs (Liang, 2013)]]

IGCC and USC are lowest on different element, but since these all have different global warming potential it is difficult to tell which is the most efficient.

[[File:GWP.PNG|thumb|border|center|middle|upright=4|link=|atl=| Quantitative environemntal measures (Liang, 2013)]]

This table breaks down the different measures of which systems are better

[[File:Measures.PNG|thumb|border|center|middle|upright=4|link=|atl=| Environmental measures and their explanations (Liang, 2013)]]

And the table above defines all the relevant terms.

[[File:Price.PNG|thumb|border|center|middle|upright=4|link=|atl=| Capital cost in $/kW of clean coal power-generation technologies (Liang, 2013)]]

This demonstrates that the capital cost per unit of energy production is very high for IGCC, but shows that Super-C and USC are very competitive.

All of the data suggests that USC is the highest in both energy efficiency and net generating efficiency while competitive in price. The IGCC while potentially the best on emissions is too much more expensive to use on a wide scale (Liang, 2013).

=Solid Waste=
U.S. federal regulations require that solid waste be classified as hazardous or nonhazardous. Conditions for a classification of hazardous include: ignitability, corrosivity, reactivity, toxicity, or posing a substantial threat to the surrounding environment and its inhabitants. Hazardous waste must be treated on- or near-site before being removed in containers. Non-hazardous waste may be landfilled or incinerated in some cases. A typical estimate of costs for waste disposal is $0.03/lb for nonhazardous solids and $0.10/lb for hazardous solids. (Seider 2009 pg 609)

=References=

*Biegler LT, Grossmann IE, Westerberg AW. ''Systematic Methods of Chemical Process Design'' Prentice-Hall: Upper Saddle River, 1997.

*Broughton, J. ''Process Utility Systems: Introduction to Design, Operation and Maintenance'' Institution of Chemical Engineers: Rugby, Warwickshire, UK, 1994.

*Clean Coal Technologies. America's Coalition for Clean Coal Electricity. http://www.americaspower.org/clean-coal-technologies-1663/

*Duke Energy Company (2013). How Do Coal Fired Plants Work? Charlotte: Duke Energy.

*"Economizers." Energy Solutions Center Inc.: Boiler Consortium. http://www.cleanboiler.org/Eff_Improve/Efficiency/Economizers.asp. Accessed February 21, 2016.

*Lehman B, Worrell E. ''Electricity Production from Natural Gas Pressure Recovery Using Expansion Turbines.'' Lawrence Berkeley National Laboratory; 2001.

*"Overview of the Clean Air Act and Air Pollution." Environmental Protection Agency. November 17, 2015. http://www.epa.gov/clean-air-act-overview. Accessed February 5, 2016.

*"Oxygen Control." Energy Solutions Center Inc.: Boiler Consortium. http://www.cleanboiler.org/Eff_Improve/Efficiency/Oxygen_Control.asp. Accessed February 21, 2016.

*Peters, Max S.; Timmerhaus, Klaus D.; West, Ronald E. (2003). " Plant Design and Economics for Chemical Engineers." McGraw Hill Higher Education.

*Ruina Xu Energy (2010). "Clean coal technology development in China." Elsevier.

*Seider, Seader, Lewin, Widagdo. (2009). ''Plant Design and Economics for Chemical Engineers, 5th Edition.'' Hoboken: Wiley.

*Seider, Seader, Lewin. (2008). ''Product and Process Design Principles, 2nd Edition.'' Hoboken: Wiley.

*Summary of the Clean Water Act. United States EPA website. http://www.epa.gov/laws-regulations/summary-clean-water-act

*Towler, G.P. and Sinnot, R. (2012). ''Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design.'' Elsevier.

*Turton R.; Bailie, R.C.; Whiting, W.B.; Shaeiwitz J.A.; Bhattacharyya D. (2012). ''Analysis, Synthesis, and Design of Chemical Processes.'' Upper Saddle River: Prentice Hall.

*G.D. Ulrich, A Guide to Chemical Engineering Process Design and Economics, Wiley: New York, 1984.

*Xiaoye Liang, Zhihua Wang, Zhijun Zhou, , Zhenyu Huang, Junhu Zhou, and Kefa Cen. (2013). "Up-to-date life cycle assessment and comparison study of clean coal power generation technologies in China" Elsevier.

Utility systems

2016-02-22T05:16:19Z

Bts511: /* Technologies and Life Cycle Comparison */

Authors: David Chen, [2014] Joshua Lee, [2015] Brett Sleyster, [2016] and Tom Aunins [2016] 

Stewards: David Chen, Jian Gong, and Fengqi You

Date Presented: January 13, 2014 /Date Revised: January 14, 2014

 


__TOC__

=Introduction=
Many chemical processes do not take place at ambient temperature or pressures. In order to reach these non-ambient conditions, utilities will have to be used to raise or lower temperatures and compress gases. (Towler, Towler/UOP) Utilities often contribute 5 to 10% of the price of a product, and may come from public or private utility companies or on-site plants. For purchased utilities, costs depend on consumption, while for company-owned utilities, there will be both capital and operating costs. They include things such as steam for heating, electricity, cooling water, refrigeration, fuels such as natural gas, wastewater treatment, waste disposal, and landfill. Steam is often the largest utility cost. Cogeneration unit can supply electricity accompanied with different steam pressures. (Seider 2009)

=Electricity=

Electricity is used to power many different kinds of equipment. It has many advantages: it is efficient (> 90%), reliable, available in a wide range of power, shaft speeds, designs, lifetimes, convenience, costs, and maintenance. It is generally used up to 200 hp, and sometimes over 10,000 Hp. In chemical process plants, the electricity demand is generally determined by the work or energy required for compression, pumping, air cooling, lights, and many other items. This electricity often times is purchased from local electricity providers, but many plants generate their own electricity via sophisticated processes.

Electricity is rarely used as a primary heat utility in large-scale chemical plants for a variety of reasons. The main disadvantages of using electricity as a heat utility are as follows (Towler 2012):

:*Heat from electricity is two to three times more expensive than heat from fuels. This is attributed to the lack of efficiency when creating heat from electricity.

:*Electrical heating units are expensive, require high maintenance, and must comply with strict safety regulations.

:*Electrical heating units are unsafe compared to steam heating units. In steam systems, the physically steam controls the temperature, whereas in electrical heating units temperature is controlled by temperature controllers, which can fail or burn out.

The use of electricity carries with it some hazards depending on the environment. Extra care must be taken when using electrically-powered equipment in areas which may have combustible fluids, vapors, or dust, and where liquids may be present. (Seider 2009 pg 606)

==Conventional Power Station==

In general, most electricity is generated from a conventional coal-fired process, whether it be on-site or purchased from a provider. Coal-fired processes have been used to create electricity throughout history, and technological advances have increased its efficiency and use worldwide. In a coal-fired steam station—much like a nuclear station—water is turned into steam, which in turn drives turbine generators to produce electricity. There are several variations on how to create energy from coal, but here are the basics of how a coal-fired process works:

[[File:coalfired.png|thumb|border|center|middle|upright=4|link=|atl=|General Coal-Fired Process Diagram (Duke Energy Company)]]

*Heat is created:
::Before the coal is burned, it is pulverized to the fineness of talcum powder. It is then mixed with hot air and blown into the firebox of the boiler. Burning in suspension, the coal/air mixture provides the most complete combustion and maximum heat possible.

*Water turns to steam:
::Highly purified water, pumped through pipes inside the boiler, is turned into steam by the heat. The steam reaches temperatures of up to 1,000 degrees Fahrenheit and pressures up to 3,500 pounds per square inch, and is piped to the turbine.

*Steam turns the turbine:
::The enormous pressure of the steam pushing against a series of giant turbine blades turns the turbine shaft. The turbine shaft is connected to the shaft of the generator, where magnets spin within wire coils to produce electricity.

*Steam is converted back to water:
::After doing its work in the turbine, the steam is drawn into a condenser, a large chamber in the basement of the power plant. In this important step, millions of gallons of cool water from a nearby source (such as a river or lake) are pumped through a network of tubes running through the condenser. The cool water in the tubes converts the steam back into water that can be used over and over again in the plant.

*Repeat:
::The cooling water is returned to its source without any contamination, and the steam water is returned to the boiler to repeat the cycle.

===Advantages & Disadvantages of Coal-Fired Energy Production===

[[File:coalfiredtable.png|thumb|border|center|middle|upright=4|link=|atl=|Advantages and Disadvantages of Coal-Fired Processes (Seider 2008)]]

==Gas-Turbine Cogeneration Process==

When generating energy on-site, many plants use a gas-turbine cogeneration process. The thermal efficiency of a gas-turbine process is in the range of 70-80% while conventional power stations, such as coal-fired processes, have a 30-40% efficiency. The lower efficiency in more conventional power stations is attributed to wasted heat in the exhaust steam in the condenser. One example of a gas-turbine process is outlined in the following figure. Figure 3.1 is a gas-turbine cogeneration process with a heat recovery steam generator (waste-heat) boiler.

[[File:cogeneration.png|thumb|border|center|middle|upright=4|link=|atl=|Gas-Turbine Cogeneration with a heat recovery steam generator boiler (Towler 2012 Fig 3.1)]]

Overall, the process illustrated is not much different from a coal-fired process. The main differences are that the cogeneration process creates both electricity and a heat utility, and cogeneration processes use natural gas instead of coal. Many of the advantages and disadvantages are similar to those of the coal-fired process, but the cogeneration has a much higher efficiency, creates heat to be used in another process, and uses a more volatile and expensive fuel. The main advantage of cogeneration over coal-fired energy production is that heat is not wasted. In coal-fired processes, heat is released and wasted during electricity generation. Cogeneration captures some, if not all of the byproduct for heat, which is an extremely useful utility that will be discussed in subsequent sections. In summary, the cogeneration plant is superior to the coal-fired process because of its higher efficiency and ability to create a useful heat utility.

Obviously any engineer would design the cogeneration plant to meet at least the energy requirement necessary for plant operation, but cogeneration plants often times are designed to exceed the plant electricity requirement to drive another source of capital. Many describe this scenario as a "make or buy" scenario (Towler 2012). This scenario provides chemical producers leverage when negotiating contracts with outsourced electricity providers and this allows plants to purchase electricity at a wholesale price. This is a huge advantage of considering on-site electricity production because electricity is needed in relatively high quantities for all chemical plants. Being able to minimize electricity costs, or even profit off of electricity production is a huge economical consideration that all plants employ.

=Process Heating=
The key objective of this section is to discuss how processes are heated. Heating utilities are necessary for proper usage of distillers, reactors, condensers, and several other integral types of equipment. More specifically, steam, fired heat, and hot oil/specialized heat transfer fluids will be discussed in the following subsections.

==Steam==
Steam is the most commonly used heat utility used in chemical plants, and as a result understanding how it is used is essential in the study of Utility systems. Steam is used both as a process fluid (feedstock, diluent to absorb heat of reaction, heating agent, and stripping agent in absorbers and adsorbers ) and utility. It can be used to drive pumps and compressors, ejectors (for producing a vacuum), and heat exchangers. As one can clearly see, steam is a versatile, and useful utility.

Here are a few advantages of using steam as opposed to other methods of process heating (Towler 2012):

:*By controlling the pressure of the steam, one can control the temperature at which the heat is released. Having a strong control over the temperature is essential in several processes. =
:*Steam is an efficient heat source because the heat of condensation of steam is very high. Meaning that there is is high output per mass of utility at a constant temperature.
:*Heat exchangers that use steam are relatively cheap because condensing steam has a high heat transfer coefficient.
:*Steam is nonflammable, nontoxic, and inert to several process fluids.

Chemical plants generally have a network of pipelines exclusively for providing steam. These networks generally have steam at a low pressure, a medium pressure, and a high pressure. The image below illustrates a basic steam system.

[[File:Steam.png|thumb|border|center|middle|upright=4|link=|atl=|(Towler 2012 Fig 3.2)]]

In the diagram above, boiler feed water at a high pressure is preheated and fed to other boilers. These other boilers superheat the steam to create a high pressure and high temperature steam stream. The steam is superheated past the dew point to account for heat loss in the pipelines. A portion of the high pressure steam is used for process heating in areas of the plant that require high temperatures. The rest of the high pressure steam is turned into medium pressure steam by valves and steam turbines. The medium pressure steam is then used to heat medium temperature processes and to form low pressure steam. The low pressure steam can be used to heat low pressure processes and it can be expanded in condensing turbines to create shaft work and energy. In summary, steam can be used for an innumerable amount of action items in a plant. High pressure, medium pressure, and low pressure steam can all be used as a heat source. Low pressure steam has utilities in creating electricity and it also has several other uses.

==Fired Heat==

In many cases, processes in a plant require a heat source stronger than high pressure and temperature steam. That is when fired heat is used, which is generally at temperatures above 523K. Streams can be heated directly in the furnace tubes or via a hot oil circuit or heat transfer fluid, which will be discussed in detail in the next section. Most fired heaters use natural gas as fuel because it burns cleaner than fuel oil. A cleaner burning fuel is always advantageous due to environmental and safety concerns. Furthermore, natural gases usually result in less wear and tear in burners and fuel lines.

Depending on the application of the fired-heater, different design specifications can be implemented to make the fired-heater as efficient as possible. The basic construction of a fired heater starts with a cylindrical chamber that is lined with with refractory bricks. The stream that is to be heated flows through tubes inside of the furnace. These tubes can be arranged in several different arrangements such as, around the walls of the furnace, or in horizontal or vertical banks (Towler 2012). The figure below illustrates the basic construction of the fired-heater and varying tube arrangements.

[[File:firedheater.png|thumb|border|center|middle|upright=4|link=|atl=|Basic Fired-Heater Diagrams (Towler 2012 Figure 19.66)]]

Fuel is burned to heat the entire furnace, and the heat transfer occurs from the combustion gases inside of the furnace across the tubes that are filled with our desired stream. The heat transfer between the tubes and the furnace is accomplished primarily via radiation. Modern designs take advantage of convective heat transfer by adding a smaller section on top of the fired-heater where the combustion gases flow over banks of tubes as seen in (c) in the figure above. Heat transfer can be further improved via convection by adding fins or pins in the combustion section.

The cost of fired heat can be calculated by the cost of fuel fired. Natural gas and heating oil are traded as commodities and prices can be found at many online trading sites or business news sites (i.e., www.cnn.money.com). Past historic prices for forecasting can be found in the Oil and Gas Journal or from the U.S. Energy Information Adminstration (www.eia.gov).

===Fuels===
Fuel is burned in utility facilities such as boilers, electricity generation, and cogeneration, and can be in solid, liquid, or gas form. It can also be burned to provide heating for a process or stream or to drive pumps and compressors. The fuel is usually burned with excess air to ensure complete combustion.

A way of quantifying the amount of heat generated is by using the heating values. Higher heating value (HHV) and the lower heating value (LHV) are used. The heating is the total heat evolved by complete combustion of a fuel with dry air with both at 60 ⁰F and the flue gas after combustion brought back down to ⁰F. If the water vapor in the flue gas is not condensed, we obtain the LHV. If the water vapor is condensed, the value of heat evolved is a bit higher, and this is the HHV.
Heating values for solids and liquids are usually on a per-mass basis, and gases on a per-volume basis.
A specified amount of heating cannot be met with the amount of fuel calculated using only the HHV. There will be heat losses, the flue gas temperature will be greater than 60 ⁰F, and water in the flue gas will typically be vapor. (Seider 608)

==Hot Oil/Specialized Heat Transfer Fluids==

Specialized heat transfer fluids and hot oil circuits are used as heat sources when steam and fired heat is not appropriate. Specialized heat transfer fluids and hot oil circuits are extremely versatile in that they can be used in the temperature range of 323K to 673K. This range however is quite variable. For hot oils, the upper temperature limit is gauged based off of the thermal decomposition of the oil and coking/fouling of heat exchanger tubes.

Hot oil circuit systems are most commonly used when the plant has several small temperature heating requirements because it is more economically sound. Rather than having several fired heaters heat each small temperature requirement, it is much more economical to have one fired heater heat the hot oil and circulate that oil through each of the process to meet all of the heat needs. Hot oil systems are also generally favored over high pressure steam in processes that involve high pressure differentials between the process stream and high pressure steam. Hot oil systems are favored in this scenario because of safety concerns. If the steam were to leak, the pressure drop could cause serious safety issues.

Mineral oils are the most commonly used heat transfer fluids, and one prominent example is Dowtherm A. Dowtherm A is a combination of 26.5 wt% diphenyl in diphenyl oxide (Towler 2012) and is extremely thermally stable. These mineral oil systems generally require high flow rates.

=Process Cooling=

==Cooling Water==
Cooling water is used to cool and/or condense streams. Cooling water is usually circulated between process heat exchangers and a cooling tower. Water is cooled during downward motion by contact with air blown upwards, which can bring the water temperature to come within ~ 5 ⁰F of air’s wet-bulb temperature.Approximately 80% of the temperature reduction is due to evaporation of the cooling water and heat transfer to the surrounding air. Water can also be cooled in spray ponds and cooling ponds. Both work by providing high area for water to exchange heat with air.
Water in cooling towers is lost through drift and blowdown, and makeup is usually 1.5 to 3% of the circulating rate.
If a large natural body of water is nearby, it can be used as a source of cooling water and discharged downstream. This water is usually filtered to remove salts and impurities that may lead to fouling, but it is not treated.

==Refrigeration==
Cooling water can usually be used to cool a stream to ~ 100 ⁰F. Air can only cool to ~ 120 ⁰F. Air may be used in places where water is scarce or more costly to transport. To cool or condense streams below 100 ⁰F, a refrigerant is typically used. Chilled brine can also be used, but is less common.

Until 1995, CFC Freon R-12 (dichlorodifuloromethane) and HCFC Freon R-22 (chlorodifuloromethane) were commonly used refrigerants. However, the chlorine atom in the molecules caused the depletion of the ozone layer. The U.S. Clean Air Act Amendments of 1990 went into effect in 1995, and the production of these refrigerants has since stopped or been greatly reduced.

Cost estimates are based on ton-day of refrigeration, where a ton is the heat that needs to be removed to freeze 1 ton per day of water at 32 ⁰F. Substitutes have since been developed. R-134a is often used in place of R-12. According to the EPA, R-134a is not combustible at ambient conditions, and is essentially non-toxic under 400 ppm, and is not ozone-depleting. (Seider pg 607)

=Energy Efficiency=
One of the chief concerns in selecting and designing process utility systems for heating and cooling is how to achieve the most energy efficient design. There are countless means by which plants lose energy, two of the foremost being through the mixing of different temperature or pressure streams and through the disposal of warmed cooling water. (Seider, 2009) Proper utilities design can help mitigate each of these losses as well as many others. The energy efficiency of a plant will depend primarily on the heating and cooling methods that are being used and the overall system design itself. These two parameters are important in determining how well energy is transferred to the desired media as well as how well that energy is recovered and recycled.

==Hot Utility Efficiency==
As mentioned above, the most commonly used utilities for process heating in large scale processes are steam, fired heat, and hot oil heaters. Of these, steam is the most commonly used. Electricity, while efficient at creating power, is not a viable source of heat in large industrial processes. Common ranges of heating efficiency for these three methods are displayed in Table 1. (Towler and Sinnott, 2012; Broughton, 1994)

{| class="wikitable"
|+Table 1: Process Heating Efficiencies
|-
! Process Heating Method
! Typical Efficiency
|-
| Steam (via coal boiler)
| 72%
|-
| Steam (via gas boiler)
| 73%
|-
| Steam (via oil boiler)
| 75%
|-
| Fired Heat w/ Convective Section
| 85%
|-
| Fired Heat w/o Convective Section
| 60%
|-
| Hot Oil Heaters/Vaporizers
| 80-85%
|}

===Steam===
As steam is so popular for heating purposes, it is useful to analyze the numerous ways in which losses can occur in steam systems. There are five primary sources of inefficiency and heat loss in the generation and distribution of steam throughout a process plant:

*The heat content of boiler exhaust gas
*Incomplete combustion of boiler fuel
*Radiant losses from the boiler exterior
*Blowdown losses
*Distribution losses (pipe transport, steam traps, etc.)

The first four of these losses take place at the boiler and contribute to the heating efficiencies seen in Table 1 for steam created with coal, gas, and oil. Several methods can be used to minimize these losses, one of the foremost being the control of air-to-fuel ratio in the boiler. This ratio must be managed by weighing losses due to uncombusted fuel against losses due to the heat content of excess exhaust gas. (Broughton, 1994)

[[File:Air-fuel ratio.PNG|thumb|border|center|middle|upright=4|link=|atl=|Air-to-fuel ratio relationship with heat losses (Broughton Fig 2.3)]]

The air-to-fuel ratio can be optimized using a feedback process controller. The control system will analyze the oxygen content of exhaust air and adjust the incoming air flow rate to achieve a set percentage of excess air. While desired excess oxygen will vary depending on the type of fuel, it is consistently seen that in the minimum loss-range a 1% increase in excess air will result in a 1% decrease in efficiency. ("Oxygen Control")

In addition to air-to-fuel ratio management, steam heat losses at the boiler can be mitigated via energy recovery, which is discussed in further detail [[Utility_systems#Energy_Recovery|below]].

Outside the boiler, losses in distribution of steam throughout facilities can decrease energy efficiency by a significant amount. This can cause up to a 60% increase in losses, but typically results in an overall steam efficiency of 50-55% (down from approximately 75% at the boiler). (Broughton, 1994) There are several ways that this issue can be addressed. First and foremost are steam metering systems, which can be used to monitor heat supply to separate sections of the process facility, check efficiency of fuel use, and determine which processes in a given plant have the highest steam requirements. Another method of minimizing transport losses is to decentralize steam generation systems. It can be advantageous to have numerous smaller boilers rather than a single centralized boiler so that steam does not have to travel as long of distances to reach its destination.

==Cold Utility Efficiency==
Efficiency in cooling processes depends more on the method used, and by extension the amount of coolant needed. Water and air utility efficiencies depend primarily on the fluid flow required to maintain the system at a desired temperature, while powered refrigeration utilities (for colder processes) have efficiencies at approximately 60%--but ranging up to 90%--of Carnot cycle efficiency, a metric of ideal refrigeration efficiency. (Towler and Sinnott, 2012) Cooling systems represent by definition a loss of energy from the main process to the utility stream, and as such it is often useful to find other uses for the heated media before discharge.

==Energy Recovery==
Recovery and recycle of energy is perhaps the most important aspect of creating an energy efficient plant design, and it is important for process engineers to fully consider possibilities for heat recovery in order to aid in economic viability.

===Process Heat Exchange===
Heat exchanger networks are a very common energy recovery method in industrial processes. These networks most frequently allow energy from heated product streams to be transferred to feed streams that must be brought up to process temperature. (Biegler, 1997) More information on the function and design of heat exchanger networks can be found on the [[Heat_exchanger#Heat_Exchanger_Networks|heat exchanger wiki page]]. The following are several examples of energy recovery via heat exchange that are used in industrial processes.

In distillation columns the bottoms and distillate effluents have the potential for energy exchange. Though the condenser at the top of the column cannot supply its waste heat to the reboiler due to their respective temperatures, the effluent streams can supply heat to the feed via a feed-effluent exchanger. This reduces the utility requirements to raise the feed to column temperature. (Biegler, 1997)

[[File:Feed_sterilization.jpg|thumb|border|right|middle|upright=4|link=|atl=|Feed sterilization schematic. (Towler and Sinnott Fig. 3.30)]]

Feed sterilization, commonly used in the food industry, is a common application for heat recovery through process stream heat exchange. In this application, the feed must be heated for a certain amount of time to kill any biological contaminants, after which it can be used to heat the new raw feed for sterilization. This reduces energy demands on the steam heater and thus reduces cost. (Towler and Sinnott, 2012)

In multi-vessel batch processes it can be advantageous to exchange heat as the process fluid is being transferred between vessels. Like the previous examples, this reduces the utility needed to bring the colder feed up to process temperature, thus reducing costs. (Towler and Sinnott, 2012)

[[File:Batch_heat_exchange.jpg|thumb|border|center|middle|upright=4|link=|atl=|Multi-vessel batch heat exchange schematic (Towler and Sinnott Fig. 3.31)]]

===Utility Regeneration===
[[File:Waste_heat_boiler.jpg|thumb|border|right|middle|200px|upright=4|link=|atl=|Industrial modular HRSG]]
When recovery of waste heat via transfer to other process streams is inconvenient or impossible, energy efficiency can still be improved through the regeneration of utilities. This is commonly done through the regeneration of steam by removing heat from exiting streams or from highly exothermic reactions. Waste heat in exiting streams can be removed via heat recovery steam generators (HRSGs), and is most often used on exiting gas streams. Heat recovery from reactions is a viable option when the reactor temperature will be at 150 C or above, as this will create steam at high enough pressure to be used in other processes. (Towler and Sinnott, 2012)

In the case of steam, waste heat and water treatment losses can be recovered from the utility generation process itself. One of the most common ways to do this is with an economizer. As seen in the schematic below, economizers have heat exchangers fit to the exhaust gas flow in order to transfer waste heat from these gases to the incoming boiler feed water. This can result in fuel energy savings of approximately 15% for typical excess air percentages in steam boilers. (Broughton, 1994; "Economizers")

[[File:Economizer.png|thumb|border|center|middle|upright=4|link=|atl=|Economizer schematic ("Economizers") and economizer energy recovery correlations (Broughton Fig 2.2)]]

[[File:Condensate recycle.PNG|thumb|border|right|middle|200px|upright=4|link=|atl=|Condensate return fuel savings (Broughton Fig 2.4)]]

Additionally, there are still more opportunities for recovery in steam generation processes. One typical case is the use of waste heat to preheat air entering the boiler. Optimistically, this can result in a 5% improvement in heat recovery for the system. Another method is the recycle of steam condensate back to the boiler. This has two main benefits: it retains the heat still present in the condensed steam and it retains the pretreatment chemicals added to the steam. (Broughton, 1994)

Furthermore, there is opportunity for energy recovery in the expansion of compressed gas through a turbine to create electricity, a process that can be economically viable given sufficiently high flows or pressure. One of the primary application of this type of energy recovery is in the creation of medium pressure and low pressure steam. In most processes, all steam is generated as high pressure steam and can be expanded through a turbine to decrease its pressure. Such technology has also been used in processes to synthesize ammonia, perform air separations, and synthesize nitric acid. (Towler and Sinnott, 2012) Recently, however, there has been a particularly strong interest for energy recovery in the natural gas industry, when gas is decompressed from major pipelines to residential low-pressure pipelines. A 2001 study estimated that there is the potential to recover 21 TWh, representing 11% of natural gas transport energy, via gas expansion. (Lehman)

=Process water and boiler-feed water=
Process water is water that will be directly used in the process. Boiler-feed water (BFW) is used to produce steam. Both may need to be purified to prevent impurities from contaminating a process or from foul equipment. It can be used as a cooling stream when the temperature of the stream to be cooled is greater than ~300 ⁰F. Cost of BFW can be partially offset by the steam credit.

Process water that undergoes moderate pretreatment can cost ~ $0.75/1,000 gal.

Extensive treatment ~ $6.00/1,000 gal.

Sterilized for pharmaceutical processes ~ $550/1,000 gal. (Seider pg 608)

=Demineralized Water=
In demineralized water, minerals have been removed by ion exchange. In boiler feed water, this is to prevent salt deposition, corrosion, formation of foam, and sluicing. In process water, the ions may contaminate the process.

=Waste Treatment=
Most chemical processes will produce some sort of waste. Disposal occurs to the atmosphere (in the case of some gases), sewers, body of water, or a landfill. Waste may require some treatment before disposal to meet regulations. Depending on process economics, byproducts may be recovered and processed. (Seider 2009 pg 609)

==Wastewater Treatment==
The United States EPA regulates industrial wastewater disposal through the Clean Water Act, introduced in 1948 as the Federal Water Pollution Control Act and amended to its current form in 1972. The sweeping 1972 amendments allowed the EPA to prevent industries and persons from discharging contaminated water into fresh water sources and set water quality standards. (Summary of the Clean Water Act) In accordance with this law, process plants in the United States treat wastewater at on-site or near-site treatment centers before releasing it into the surrounding environment.

Wastewater effluent streams, along with water runoff from around the plant, are treated to control for pH, toxicity, suspended solids, and biological oxygen demand (for aquatic life protection) prior to discharge. Each of these controls is typically addressed with a separate method. Acidity and basicity is balanced through the addition of an acid or alkaline solution. Toxic wastewater may be treated with chemical processes or simply diluted to safe concentrations. Suspended solids can be removed via filtration and/or with clarifiers. Oxygen demand of wastewater can be mitigated using activated sludge treatment processes. Once the water quality complies with the EPA, and state-mandated, regulations, it can be safely released. More information on the large number of industry-specific guidelines for waste effluent can be found on the EPA website (http://www.epa.gov/eg/industrial-effluent-guidelines).

=Air-Pollution Management=
===Introduction===

In the United States air pollution is regulated in the Clean Air Act, and almost all pollutant emitting plants are regulated under this law. The types of plants that can release significant emissions include petroleum refineries, sulfur recovery plants, carbon-black plants, fuel conversion plants, chemical process plants, fossil fuel plants, and petroleum storage and transfer facilities. To receive permission to construct a plant must undergo a review to show that it will not cause a violation of the Ambient Air Quality Standards(Peters, 2003).

==Methods==

There are two major types of pollutants that are released into the air, particulates and and gaseous pollutants. Particulates can be removed with mechanical forces while gaseous pollutants typically need to removed by chemical or physical means (Peters, 2003).

Particulates and volatile pollutants that need to be removed before disposal may be present. Particle removal equipment includes: cyclones, wet scrubbers, electrostatic precipitators, and fabric-filter systems (Seider 2009) such as bag filters (. The two charts below are from ''Plant Design and Economics for Chemical Engineers'' and show the types of equipment, separation methods, and particle sizes in different pollutant separation technologies (Peters, 1991).

[[File:AirPolChart.png|thumb|border|center|middle|upright=4|link=|atl=|Different Methods of Particle Separation, the Particle Sizes they Can Remove, and the Technologies Used]]

[[File:AirPolTable.png|thumb|border|center|middle|upright=4|link=|atl=| This Chart Shows the Specifications and Limitations of Different Separations Technologies Including Particle Size, Efficiencies, and Temperatures(Peters, 1991)]]

Methods for removing inorganic and organic gaseous pollutants include: absorption, adsorption, condensation, and combustion (Seider 2009 pg 609). A list of typical of gases pollutants and their sources from ''Plant Design and Economics for Chemical Engineers'' is shown below (Peters, 1991).

[[File:GaseousPollutants.png|thumb|border|center|middle|upright=4|link=|atl=| Common Gaseous Pollutants and their Sources]]

Typically Gas-liquid absorption processes are done completed in a vertical, countercurrent, flow through packed, plate, or spray towers. These systems require good liquid-gas contact and proper equipment. These systems also often have significant energy consumption because of large pressure drops (Peters, 2003). For high volume systems absorption by scrubbing with water or another solvent is the most widely used method (Towler, 2012). Dry adsorbents can be used to remove the last races of gaseous pollutants. Adsorption typically requires blowers, condensers, separators, and controls. You also typically need two packed beds so that one can be used while the other is regenerated. Examples of adsorbents are molecular sieves and activated carbon. Incineration is typically used when there are gas streams that have no recovery value. This can be done with direct flame or catalytic oxidation. Catalytic oxidation usually has higher capital costs, but lower operating costs because it does not require fuel.
===Wet Scrubbers===
Wet scrubber use lime or limestone and water to remove SO2 and acid gases. The mixture can be injected into a scrubber or the gases can be bubbled through this mixture. This results in removal of 90-98% of SO2 and and acid gases (Clean Coal Technologies).

===Dry Scrubbers===
Dry scrubber blow powdered adsorbents into a vessel with gases and then after it has captured the SO2 and acid gases it is separated from gas using a fabric filter. These systems remove 90-93% of the contaminants (Clean Coal Technologies).

===Low-NOx Burners===
The purpose of low NOx burners is to decrease the amount of NOx created when the coal is burned. This is done by injecting coal and air in boilers. This can result in 40-50% NOx reduction. If air is injected into the area above the burner this can actually cause almost 70% NOx reduction (Clean Coal Technologies).

===Selective Non-Catalytic and Catalytic Reductions===
These systems inject ammonia into gases to remove NOx. The Catalytic reductions add a catalyst to the ammonia being injected to aid in the removal. Con-catalytic reductions result in about 35% removal, but adding a catalyst can increase that amount to about 90%. The catalytic-reduction can remove up to 80% of mercury as well (Clean Coal Technologies).

===Fabric Filters===
Also known as baghouses,these [[Separation processes#Filtration| filters]] remove particulates by passing air through filters. These can separate as much as 99.9% of particulate matter (Clean Coal Technologies).

===Electrostatic Precipitators===
Electrostatic Precicpitators remove particulate matter as gas passes through a device that has charged metal plates. The particles are then removed because of static electricity. These systems can remove between 99-99.9% of particulate matter (Clean Coal Technologies).

===Super-critical Boilers===
Super-Critical Boilers and Ultra-Supercritical Boilers operate at temperature and pressure higher than regular boilers. By operating at higher temperature these systems become more efficient. Super Critical Boilers typically have 10%-20% CO2 emissions than other similar sub-critical technologies. Ultra-Supercritical boilers can be as much as 30% more efficent than sub-critical technologies (Clean Coal Technologies).

===More Information===

For more information regarding some of these separations equipment see [[Solids-involved equipment]]

For more information Regarding Cyclones see [[Separation processes#Cyclones| Cyclones]], and for modeling cyclones see their[[Solids-involved equipment#HYSYS Simulation |HYSYS Simulation]]

You can find more information on [[Separation processes#Absorption| Absorption]] and [[Separation processes#Adsorption| Adsorption]] in [[Separation processes]]

===Outcomes===

====United States Common Pollutant Emissions====

The United States implemented the Clean Air Act in 1970 and since then emissions in the U.S. have been drastically reduced. As a matter of fact despite increases in the population in the last 35 years the amount of pollutants emitted have decreased by almost 70% (EPA). More importantly this demonstrates that reasonable efforts can be put towards environmental protection without causing too much harm to industry.

[[File:EPA.png|thumb|border|center|middle|upright=4|link=|atl=| GDP and Other Growth Factors Vs Common Pollutants Released]]

===China Clean Coal===

====Success Through 2005====

Coal is a very inexpensive and abundant source of energy and is abundant in China (Xu 2010). In China Coal the cause of 90% of SO2 emissions, 70% of dust emissions, and 67% of NOx emissions, and 70% of CO2 emissions. These numbers are staggering considering that Coal has never been more than 50 percent of China's Energy Supply (Xu 2010).

[[File:Coal.jpg|thumb|border|center|middle|upright=4|link=|atl=| Coal Consumption for Power and the Percentage of Coal Used (Xu 2010)]]

Despite the increasing coal consumption, high efficiency electric dust removal systems with efficiencies that are as high as 99.6% have greatly decreased soot emissions to 32% below 1980's levels as of 2005.

[[File:Soot.jpg|thumb|border|center|middle|upright=4|link=|atl=| SO2 and Soot Emissions from 1981 to 2005 (Xu 2010)]]

====Technologies and Life Cycle Comparison====
As of 2010 China was consuming 48.2% of coal globally. Four potential technologies integrated gasification combined cycle (IGCC), sub-critical coal power generation (Sub-C); super-critical coal power generation (Super-C) ultrasuper-critical coal power generation (USC). These technologies are compared mostly on the basis of net generating efficiency and efficiency. Net generating efficiency is the output of the plant divided by the total available energy in the fuel used (Liang, 2013).

where efficiency is define as: η=E/Etotal

and

Etotal=Emining+Etransportation+Egeneration

[[File:technologies.PNG|thumb|border|center|middle|upright=4|link=|atl=|A comparison of many of the parameters important to lfe-cycle analysis (Liang, 2013)]]

This demonstrates that USC has the best net generation efficiency, while also having the largest capacity for a single system.

[[File:Lifecycle2.PNG|thumb|border|center|middle|upright=4|link=|atl=| (Liang, 2013)]]

This breaks down the energy efficiency of each system and demonstrates that USC is the most efficient

[[File:Lifecycle.PNG|thumb|border|center|middle|upright=4|link=|atl=| Various pollutant outputs (Liang, 2013)]]

IGCC and USC are lowest on different element, but since these all have different global warming potential it is difficult to tell which is the most efficient.

[[File:GWP.PNG|thumb|border|center|middle|upright=4|link=|atl=| Environemntal measures and their explanations (Liang, 2013)]]

This table breaks down the different measures of which systems are better

[[File:Measures.PNG|thumb|border|center|middle|upright=4|link=|atl=| Environemntal measures and their explanations (Liang, 2013)]]

And the table above defines all the relevant terms.

[[File:Price.PNG|thumb|border|center|middle|upright=4|link=|atl=| Capital cost in $/kW of clean coal power-generation technologies (Liang, 2013)]]

This demonstrates that the capital cost per unit of energy production is very high for IGCC, but shows that Super-C and USC are very competitive.

All of the data suggests that USC is the highest in both energy efficiency and net generating efficiency while competitive in price. The IGCC while potentially the best on emissions is too much more expensive to use on a wide scale (Liang, 2013).

=Solid Waste=
U.S. federal regulations require that solid waste be classified as hazardous or nonhazardous. Conditions for a classification of hazardous include: ignitability, corrosivity, reactivity, toxicity, or posing a substantial threat to the surrounding environment and its inhabitants. Hazardous waste must be treated on- or near-site before being removed in containers. Non-hazardous waste may be landfilled or incinerated in some cases. A typical estimate of costs for waste disposal is $0.03/lb for nonhazardous solids and $0.10/lb for hazardous solids. (Seider 2009 pg 609)

=References=

*Biegler LT, Grossmann IE, Westerberg AW. ''Systematic Methods of Chemical Process Design'' Prentice-Hall: Upper Saddle River, 1997.

*Broughton, J. ''Process Utility Systems: Introduction to Design, Operation and Maintenance'' Institution of Chemical Engineers: Rugby, Warwickshire, UK, 1994.

*Clean Coal Technologies. America's Coalition for Clean Coal Electricity. http://www.americaspower.org/clean-coal-technologies-1663/

*Duke Energy Company (2013). How Do Coal Fired Plants Work? Charlotte: Duke Energy.

*"Economizers." Energy Solutions Center Inc.: Boiler Consortium. http://www.cleanboiler.org/Eff_Improve/Efficiency/Economizers.asp. Accessed February 21, 2016.

*Lehman B, Worrell E. ''Electricity Production from Natural Gas Pressure Recovery Using Expansion Turbines.'' Lawrence Berkeley National Laboratory; 2001.

*"Overview of the Clean Air Act and Air Pollution." Environmental Protection Agency. November 17, 2015. http://www.epa.gov/clean-air-act-overview. Accessed February 5, 2016.

*"Oxygen Control." Energy Solutions Center Inc.: Boiler Consortium. http://www.cleanboiler.org/Eff_Improve/Efficiency/Oxygen_Control.asp. Accessed February 21, 2016.

*Peters, Max S.; Timmerhaus, Klaus D.; West, Ronald E. (2003). " Plant Design and Economics for Chemical Engineers." McGraw Hill Higher Education.

*Ruina Xu Energy (2010). "Clean coal technology development in China." Elsevier.

*Seider, Seader, Lewin, Widagdo. (2009). ''Plant Design and Economics for Chemical Engineers, 5th Edition.'' Hoboken: Wiley.

*Seider, Seader, Lewin. (2008). ''Product and Process Design Principles, 2nd Edition.'' Hoboken: Wiley.

*Summary of the Clean Water Act. United States EPA website. http://www.epa.gov/laws-regulations/summary-clean-water-act

*Towler, G.P. and Sinnot, R. (2012). ''Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design.'' Elsevier.

*Turton R.; Bailie, R.C.; Whiting, W.B.; Shaeiwitz J.A.; Bhattacharyya D. (2012). ''Analysis, Synthesis, and Design of Chemical Processes.'' Upper Saddle River: Prentice Hall.

*G.D. Ulrich, A Guide to Chemical Engineering Process Design and Economics, Wiley: New York, 1984.

*Xiaoye Liang, Zhihua Wang, Zhijun Zhou, , Zhenyu Huang, Junhu Zhou, and Kefa Cen. (2013). "Up-to-date life cycle assessment and comparison study of clean coal power generation technologies in China" Elsevier.

Utility systems

2016-02-22T05:14:19Z

Bts511: /* More Information */

Authors: David Chen, [2014] Joshua Lee, [2015] Brett Sleyster, [2016] and Tom Aunins [2016] 

Stewards: David Chen, Jian Gong, and Fengqi You

Date Presented: January 13, 2014 /Date Revised: January 14, 2014

 


__TOC__

=Introduction=
Many chemical processes do not take place at ambient temperature or pressures. In order to reach these non-ambient conditions, utilities will have to be used to raise or lower temperatures and compress gases. (Towler, Towler/UOP) Utilities often contribute 5 to 10% of the price of a product, and may come from public or private utility companies or on-site plants. For purchased utilities, costs depend on consumption, while for company-owned utilities, there will be both capital and operating costs. They include things such as steam for heating, electricity, cooling water, refrigeration, fuels such as natural gas, wastewater treatment, waste disposal, and landfill. Steam is often the largest utility cost. Cogeneration unit can supply electricity accompanied with different steam pressures. (Seider 2009)

=Electricity=

Electricity is used to power many different kinds of equipment. It has many advantages: it is efficient (> 90%), reliable, available in a wide range of power, shaft speeds, designs, lifetimes, convenience, costs, and maintenance. It is generally used up to 200 hp, and sometimes over 10,000 Hp. In chemical process plants, the electricity demand is generally determined by the work or energy required for compression, pumping, air cooling, lights, and many other items. This electricity often times is purchased from local electricity providers, but many plants generate their own electricity via sophisticated processes.

Electricity is rarely used as a primary heat utility in large-scale chemical plants for a variety of reasons. The main disadvantages of using electricity as a heat utility are as follows (Towler 2012):

:*Heat from electricity is two to three times more expensive than heat from fuels. This is attributed to the lack of efficiency when creating heat from electricity.

:*Electrical heating units are expensive, require high maintenance, and must comply with strict safety regulations.

:*Electrical heating units are unsafe compared to steam heating units. In steam systems, the physically steam controls the temperature, whereas in electrical heating units temperature is controlled by temperature controllers, which can fail or burn out.

The use of electricity carries with it some hazards depending on the environment. Extra care must be taken when using electrically-powered equipment in areas which may have combustible fluids, vapors, or dust, and where liquids may be present. (Seider 2009 pg 606)

==Conventional Power Station==

In general, most electricity is generated from a conventional coal-fired process, whether it be on-site or purchased from a provider. Coal-fired processes have been used to create electricity throughout history, and technological advances have increased its efficiency and use worldwide. In a coal-fired steam station—much like a nuclear station—water is turned into steam, which in turn drives turbine generators to produce electricity. There are several variations on how to create energy from coal, but here are the basics of how a coal-fired process works:

[[File:coalfired.png|thumb|border|center|middle|upright=4|link=|atl=|General Coal-Fired Process Diagram (Duke Energy Company)]]

*Heat is created:
::Before the coal is burned, it is pulverized to the fineness of talcum powder. It is then mixed with hot air and blown into the firebox of the boiler. Burning in suspension, the coal/air mixture provides the most complete combustion and maximum heat possible.

*Water turns to steam:
::Highly purified water, pumped through pipes inside the boiler, is turned into steam by the heat. The steam reaches temperatures of up to 1,000 degrees Fahrenheit and pressures up to 3,500 pounds per square inch, and is piped to the turbine.

*Steam turns the turbine:
::The enormous pressure of the steam pushing against a series of giant turbine blades turns the turbine shaft. The turbine shaft is connected to the shaft of the generator, where magnets spin within wire coils to produce electricity.

*Steam is converted back to water:
::After doing its work in the turbine, the steam is drawn into a condenser, a large chamber in the basement of the power plant. In this important step, millions of gallons of cool water from a nearby source (such as a river or lake) are pumped through a network of tubes running through the condenser. The cool water in the tubes converts the steam back into water that can be used over and over again in the plant.

*Repeat:
::The cooling water is returned to its source without any contamination, and the steam water is returned to the boiler to repeat the cycle.

===Advantages & Disadvantages of Coal-Fired Energy Production===

[[File:coalfiredtable.png|thumb|border|center|middle|upright=4|link=|atl=|Advantages and Disadvantages of Coal-Fired Processes (Seider 2008)]]

==Gas-Turbine Cogeneration Process==

When generating energy on-site, many plants use a gas-turbine cogeneration process. The thermal efficiency of a gas-turbine process is in the range of 70-80% while conventional power stations, such as coal-fired processes, have a 30-40% efficiency. The lower efficiency in more conventional power stations is attributed to wasted heat in the exhaust steam in the condenser. One example of a gas-turbine process is outlined in the following figure. Figure 3.1 is a gas-turbine cogeneration process with a heat recovery steam generator (waste-heat) boiler.

[[File:cogeneration.png|thumb|border|center|middle|upright=4|link=|atl=|Gas-Turbine Cogeneration with a heat recovery steam generator boiler (Towler 2012 Fig 3.1)]]

Overall, the process illustrated is not much different from a coal-fired process. The main differences are that the cogeneration process creates both electricity and a heat utility, and cogeneration processes use natural gas instead of coal. Many of the advantages and disadvantages are similar to those of the coal-fired process, but the cogeneration has a much higher efficiency, creates heat to be used in another process, and uses a more volatile and expensive fuel. The main advantage of cogeneration over coal-fired energy production is that heat is not wasted. In coal-fired processes, heat is released and wasted during electricity generation. Cogeneration captures some, if not all of the byproduct for heat, which is an extremely useful utility that will be discussed in subsequent sections. In summary, the cogeneration plant is superior to the coal-fired process because of its higher efficiency and ability to create a useful heat utility.

Obviously any engineer would design the cogeneration plant to meet at least the energy requirement necessary for plant operation, but cogeneration plants often times are designed to exceed the plant electricity requirement to drive another source of capital. Many describe this scenario as a "make or buy" scenario (Towler 2012). This scenario provides chemical producers leverage when negotiating contracts with outsourced electricity providers and this allows plants to purchase electricity at a wholesale price. This is a huge advantage of considering on-site electricity production because electricity is needed in relatively high quantities for all chemical plants. Being able to minimize electricity costs, or even profit off of electricity production is a huge economical consideration that all plants employ.

=Process Heating=
The key objective of this section is to discuss how processes are heated. Heating utilities are necessary for proper usage of distillers, reactors, condensers, and several other integral types of equipment. More specifically, steam, fired heat, and hot oil/specialized heat transfer fluids will be discussed in the following subsections.

==Steam==
Steam is the most commonly used heat utility used in chemical plants, and as a result understanding how it is used is essential in the study of Utility systems. Steam is used both as a process fluid (feedstock, diluent to absorb heat of reaction, heating agent, and stripping agent in absorbers and adsorbers ) and utility. It can be used to drive pumps and compressors, ejectors (for producing a vacuum), and heat exchangers. As one can clearly see, steam is a versatile, and useful utility.

Here are a few advantages of using steam as opposed to other methods of process heating (Towler 2012):

:*By controlling the pressure of the steam, one can control the temperature at which the heat is released. Having a strong control over the temperature is essential in several processes. =
:*Steam is an efficient heat source because the heat of condensation of steam is very high. Meaning that there is is high output per mass of utility at a constant temperature.
:*Heat exchangers that use steam are relatively cheap because condensing steam has a high heat transfer coefficient.
:*Steam is nonflammable, nontoxic, and inert to several process fluids.

Chemical plants generally have a network of pipelines exclusively for providing steam. These networks generally have steam at a low pressure, a medium pressure, and a high pressure. The image below illustrates a basic steam system.

[[File:Steam.png|thumb|border|center|middle|upright=4|link=|atl=|(Towler 2012 Fig 3.2)]]

In the diagram above, boiler feed water at a high pressure is preheated and fed to other boilers. These other boilers superheat the steam to create a high pressure and high temperature steam stream. The steam is superheated past the dew point to account for heat loss in the pipelines. A portion of the high pressure steam is used for process heating in areas of the plant that require high temperatures. The rest of the high pressure steam is turned into medium pressure steam by valves and steam turbines. The medium pressure steam is then used to heat medium temperature processes and to form low pressure steam. The low pressure steam can be used to heat low pressure processes and it can be expanded in condensing turbines to create shaft work and energy. In summary, steam can be used for an innumerable amount of action items in a plant. High pressure, medium pressure, and low pressure steam can all be used as a heat source. Low pressure steam has utilities in creating electricity and it also has several other uses.

==Fired Heat==

In many cases, processes in a plant require a heat source stronger than high pressure and temperature steam. That is when fired heat is used, which is generally at temperatures above 523K. Streams can be heated directly in the furnace tubes or via a hot oil circuit or heat transfer fluid, which will be discussed in detail in the next section. Most fired heaters use natural gas as fuel because it burns cleaner than fuel oil. A cleaner burning fuel is always advantageous due to environmental and safety concerns. Furthermore, natural gases usually result in less wear and tear in burners and fuel lines.

Depending on the application of the fired-heater, different design specifications can be implemented to make the fired-heater as efficient as possible. The basic construction of a fired heater starts with a cylindrical chamber that is lined with with refractory bricks. The stream that is to be heated flows through tubes inside of the furnace. These tubes can be arranged in several different arrangements such as, around the walls of the furnace, or in horizontal or vertical banks (Towler 2012). The figure below illustrates the basic construction of the fired-heater and varying tube arrangements.

[[File:firedheater.png|thumb|border|center|middle|upright=4|link=|atl=|Basic Fired-Heater Diagrams (Towler 2012 Figure 19.66)]]

Fuel is burned to heat the entire furnace, and the heat transfer occurs from the combustion gases inside of the furnace across the tubes that are filled with our desired stream. The heat transfer between the tubes and the furnace is accomplished primarily via radiation. Modern designs take advantage of convective heat transfer by adding a smaller section on top of the fired-heater where the combustion gases flow over banks of tubes as seen in (c) in the figure above. Heat transfer can be further improved via convection by adding fins or pins in the combustion section.

The cost of fired heat can be calculated by the cost of fuel fired. Natural gas and heating oil are traded as commodities and prices can be found at many online trading sites or business news sites (i.e., www.cnn.money.com). Past historic prices for forecasting can be found in the Oil and Gas Journal or from the U.S. Energy Information Adminstration (www.eia.gov).

===Fuels===
Fuel is burned in utility facilities such as boilers, electricity generation, and cogeneration, and can be in solid, liquid, or gas form. It can also be burned to provide heating for a process or stream or to drive pumps and compressors. The fuel is usually burned with excess air to ensure complete combustion.

A way of quantifying the amount of heat generated is by using the heating values. Higher heating value (HHV) and the lower heating value (LHV) are used. The heating is the total heat evolved by complete combustion of a fuel with dry air with both at 60 ⁰F and the flue gas after combustion brought back down to ⁰F. If the water vapor in the flue gas is not condensed, we obtain the LHV. If the water vapor is condensed, the value of heat evolved is a bit higher, and this is the HHV.
Heating values for solids and liquids are usually on a per-mass basis, and gases on a per-volume basis.
A specified amount of heating cannot be met with the amount of fuel calculated using only the HHV. There will be heat losses, the flue gas temperature will be greater than 60 ⁰F, and water in the flue gas will typically be vapor. (Seider 608)

==Hot Oil/Specialized Heat Transfer Fluids==

Specialized heat transfer fluids and hot oil circuits are used as heat sources when steam and fired heat is not appropriate. Specialized heat transfer fluids and hot oil circuits are extremely versatile in that they can be used in the temperature range of 323K to 673K. This range however is quite variable. For hot oils, the upper temperature limit is gauged based off of the thermal decomposition of the oil and coking/fouling of heat exchanger tubes.

Hot oil circuit systems are most commonly used when the plant has several small temperature heating requirements because it is more economically sound. Rather than having several fired heaters heat each small temperature requirement, it is much more economical to have one fired heater heat the hot oil and circulate that oil through each of the process to meet all of the heat needs. Hot oil systems are also generally favored over high pressure steam in processes that involve high pressure differentials between the process stream and high pressure steam. Hot oil systems are favored in this scenario because of safety concerns. If the steam were to leak, the pressure drop could cause serious safety issues.

Mineral oils are the most commonly used heat transfer fluids, and one prominent example is Dowtherm A. Dowtherm A is a combination of 26.5 wt% diphenyl in diphenyl oxide (Towler 2012) and is extremely thermally stable. These mineral oil systems generally require high flow rates.

=Process Cooling=

==Cooling Water==
Cooling water is used to cool and/or condense streams. Cooling water is usually circulated between process heat exchangers and a cooling tower. Water is cooled during downward motion by contact with air blown upwards, which can bring the water temperature to come within ~ 5 ⁰F of air’s wet-bulb temperature.Approximately 80% of the temperature reduction is due to evaporation of the cooling water and heat transfer to the surrounding air. Water can also be cooled in spray ponds and cooling ponds. Both work by providing high area for water to exchange heat with air.
Water in cooling towers is lost through drift and blowdown, and makeup is usually 1.5 to 3% of the circulating rate.
If a large natural body of water is nearby, it can be used as a source of cooling water and discharged downstream. This water is usually filtered to remove salts and impurities that may lead to fouling, but it is not treated.

==Refrigeration==
Cooling water can usually be used to cool a stream to ~ 100 ⁰F. Air can only cool to ~ 120 ⁰F. Air may be used in places where water is scarce or more costly to transport. To cool or condense streams below 100 ⁰F, a refrigerant is typically used. Chilled brine can also be used, but is less common.

Until 1995, CFC Freon R-12 (dichlorodifuloromethane) and HCFC Freon R-22 (chlorodifuloromethane) were commonly used refrigerants. However, the chlorine atom in the molecules caused the depletion of the ozone layer. The U.S. Clean Air Act Amendments of 1990 went into effect in 1995, and the production of these refrigerants has since stopped or been greatly reduced.

Cost estimates are based on ton-day of refrigeration, where a ton is the heat that needs to be removed to freeze 1 ton per day of water at 32 ⁰F. Substitutes have since been developed. R-134a is often used in place of R-12. According to the EPA, R-134a is not combustible at ambient conditions, and is essentially non-toxic under 400 ppm, and is not ozone-depleting. (Seider pg 607)

=Energy Efficiency=
One of the chief concerns in selecting and designing process utility systems for heating and cooling is how to achieve the most energy efficient design. There are countless means by which plants lose energy, two of the foremost being through the mixing of different temperature or pressure streams and through the disposal of warmed cooling water. (Seider, 2009) Proper utilities design can help mitigate each of these losses as well as many others. The energy efficiency of a plant will depend primarily on the heating and cooling methods that are being used and the overall system design itself. These two parameters are important in determining how well energy is transferred to the desired media as well as how well that energy is recovered and recycled.

==Hot Utility Efficiency==
As mentioned above, the most commonly used utilities for process heating in large scale processes are steam, fired heat, and hot oil heaters. Of these, steam is the most commonly used. Electricity, while efficient at creating power, is not a viable source of heat in large industrial processes. Common ranges of heating efficiency for these three methods are displayed in Table 1. (Towler and Sinnott, 2012; Broughton, 1994)

{| class="wikitable"
|+Table 1: Process Heating Efficiencies
|-
! Process Heating Method
! Typical Efficiency
|-
| Steam (via coal boiler)
| 72%
|-
| Steam (via gas boiler)
| 73%
|-
| Steam (via oil boiler)
| 75%
|-
| Fired Heat w/ Convective Section
| 85%
|-
| Fired Heat w/o Convective Section
| 60%
|-
| Hot Oil Heaters/Vaporizers
| 80-85%
|}

===Steam===
As steam is so popular for heating purposes, it is useful to analyze the numerous ways in which losses can occur in steam systems. There are five primary sources of inefficiency and heat loss in the generation and distribution of steam throughout a process plant:

*The heat content of boiler exhaust gas
*Incomplete combustion of boiler fuel
*Radiant losses from the boiler exterior
*Blowdown losses
*Distribution losses (pipe transport, steam traps, etc.)

The first four of these losses take place at the boiler and contribute to the heating efficiencies seen in Table 1 for steam created with coal, gas, and oil. Several methods can be used to minimize these losses, one of the foremost being the control of air-to-fuel ratio in the boiler. This ratio must be managed by weighing losses due to uncombusted fuel against losses due to the heat content of excess exhaust gas. (Broughton, 1994)

[[File:Air-fuel ratio.PNG|thumb|border|center|middle|upright=4|link=|atl=|Air-to-fuel ratio relationship with heat losses (Broughton Fig 2.3)]]

The air-to-fuel ratio can be optimized using a feedback process controller. The control system will analyze the oxygen content of exhaust air and adjust the incoming air flow rate to achieve a set percentage of excess air. While desired excess oxygen will vary depending on the type of fuel, it is consistently seen that in the minimum loss-range a 1% increase in excess air will result in a 1% decrease in efficiency. ("Oxygen Control")

In addition to air-to-fuel ratio management, steam heat losses at the boiler can be mitigated via energy recovery, which is discussed in further detail [[Utility_systems#Energy_Recovery|below]].

Outside the boiler, losses in distribution of steam throughout facilities can decrease energy efficiency by a significant amount. This can cause up to a 60% increase in losses, but typically results in an overall steam efficiency of 50-55% (down from approximately 75% at the boiler). (Broughton, 1994) There are several ways that this issue can be addressed. First and foremost are steam metering systems, which can be used to monitor heat supply to separate sections of the process facility, check efficiency of fuel use, and determine which processes in a given plant have the highest steam requirements. Another method of minimizing transport losses is to decentralize steam generation systems. It can be advantageous to have numerous smaller boilers rather than a single centralized boiler so that steam does not have to travel as long of distances to reach its destination.

==Cold Utility Efficiency==
Efficiency in cooling processes depends more on the method used, and by extension the amount of coolant needed. Water and air utility efficiencies depend primarily on the fluid flow required to maintain the system at a desired temperature, while powered refrigeration utilities (for colder processes) have efficiencies at approximately 60%--but ranging up to 90%--of Carnot cycle efficiency, a metric of ideal refrigeration efficiency. (Towler and Sinnott, 2012) Cooling systems represent by definition a loss of energy from the main process to the utility stream, and as such it is often useful to find other uses for the heated media before discharge.

==Energy Recovery==
Recovery and recycle of energy is perhaps the most important aspect of creating an energy efficient plant design, and it is important for process engineers to fully consider possibilities for heat recovery in order to aid in economic viability.

===Process Heat Exchange===
Heat exchanger networks are a very common energy recovery method in industrial processes. These networks most frequently allow energy from heated product streams to be transferred to feed streams that must be brought up to process temperature. (Biegler, 1997) More information on the function and design of heat exchanger networks can be found on the [[Heat_exchanger#Heat_Exchanger_Networks|heat exchanger wiki page]]. The following are several examples of energy recovery via heat exchange that are used in industrial processes.

In distillation columns the bottoms and distillate effluents have the potential for energy exchange. Though the condenser at the top of the column cannot supply its waste heat to the reboiler due to their respective temperatures, the effluent streams can supply heat to the feed via a feed-effluent exchanger. This reduces the utility requirements to raise the feed to column temperature. (Biegler, 1997)

[[File:Feed_sterilization.jpg|thumb|border|right|middle|upright=4|link=|atl=|Feed sterilization schematic. (Towler and Sinnott Fig. 3.30)]]

Feed sterilization, commonly used in the food industry, is a common application for heat recovery through process stream heat exchange. In this application, the feed must be heated for a certain amount of time to kill any biological contaminants, after which it can be used to heat the new raw feed for sterilization. This reduces energy demands on the steam heater and thus reduces cost. (Towler and Sinnott, 2012)

In multi-vessel batch processes it can be advantageous to exchange heat as the process fluid is being transferred between vessels. Like the previous examples, this reduces the utility needed to bring the colder feed up to process temperature, thus reducing costs. (Towler and Sinnott, 2012)

[[File:Batch_heat_exchange.jpg|thumb|border|center|middle|upright=4|link=|atl=|Multi-vessel batch heat exchange schematic (Towler and Sinnott Fig. 3.31)]]

===Utility Regeneration===
[[File:Waste_heat_boiler.jpg|thumb|border|right|middle|200px|upright=4|link=|atl=|Industrial modular HRSG]]
When recovery of waste heat via transfer to other process streams is inconvenient or impossible, energy efficiency can still be improved through the regeneration of utilities. This is commonly done through the regeneration of steam by removing heat from exiting streams or from highly exothermic reactions. Waste heat in exiting streams can be removed via heat recovery steam generators (HRSGs), and is most often used on exiting gas streams. Heat recovery from reactions is a viable option when the reactor temperature will be at 150 C or above, as this will create steam at high enough pressure to be used in other processes. (Towler and Sinnott, 2012)

In the case of steam, waste heat and water treatment losses can be recovered from the utility generation process itself. One of the most common ways to do this is with an economizer. As seen in the schematic below, economizers have heat exchangers fit to the exhaust gas flow in order to transfer waste heat from these gases to the incoming boiler feed water. This can result in fuel energy savings of approximately 15% for typical excess air percentages in steam boilers. (Broughton, 1994; "Economizers")

[[File:Economizer.png|thumb|border|center|middle|upright=4|link=|atl=|Economizer schematic ("Economizers") and economizer energy recovery correlations (Broughton Fig 2.2)]]

[[File:Condensate recycle.PNG|thumb|border|right|middle|200px|upright=4|link=|atl=|Condensate return fuel savings (Broughton Fig 2.4)]]

Additionally, there are still more opportunities for recovery in steam generation processes. One typical case is the use of waste heat to preheat air entering the boiler. Optimistically, this can result in a 5% improvement in heat recovery for the system. Another method is the recycle of steam condensate back to the boiler. This has two main benefits: it retains the heat still present in the condensed steam and it retains the pretreatment chemicals added to the steam. (Broughton, 1994)

Furthermore, there is opportunity for energy recovery in the expansion of compressed gas through a turbine to create electricity, a process that can be economically viable given sufficiently high flows or pressure. One of the primary application of this type of energy recovery is in the creation of medium pressure and low pressure steam. In most processes, all steam is generated as high pressure steam and can be expanded through a turbine to decrease its pressure. Such technology has also been used in processes to synthesize ammonia, perform air separations, and synthesize nitric acid. (Towler and Sinnott, 2012) Recently, however, there has been a particularly strong interest for energy recovery in the natural gas industry, when gas is decompressed from major pipelines to residential low-pressure pipelines. A 2001 study estimated that there is the potential to recover 21 TWh, representing 11% of natural gas transport energy, via gas expansion. (Lehman)

=Process water and boiler-feed water=
Process water is water that will be directly used in the process. Boiler-feed water (BFW) is used to produce steam. Both may need to be purified to prevent impurities from contaminating a process or from foul equipment. It can be used as a cooling stream when the temperature of the stream to be cooled is greater than ~300 ⁰F. Cost of BFW can be partially offset by the steam credit.

Process water that undergoes moderate pretreatment can cost ~ $0.75/1,000 gal.

Extensive treatment ~ $6.00/1,000 gal.

Sterilized for pharmaceutical processes ~ $550/1,000 gal. (Seider pg 608)

=Demineralized Water=
In demineralized water, minerals have been removed by ion exchange. In boiler feed water, this is to prevent salt deposition, corrosion, formation of foam, and sluicing. In process water, the ions may contaminate the process.

=Waste Treatment=
Most chemical processes will produce some sort of waste. Disposal occurs to the atmosphere (in the case of some gases), sewers, body of water, or a landfill. Waste may require some treatment before disposal to meet regulations. Depending on process economics, byproducts may be recovered and processed. (Seider 2009 pg 609)

==Wastewater Treatment==
The United States EPA regulates industrial wastewater disposal through the Clean Water Act, introduced in 1948 as the Federal Water Pollution Control Act and amended to its current form in 1972. The sweeping 1972 amendments allowed the EPA to prevent industries and persons from discharging contaminated water into fresh water sources and set water quality standards. (Summary of the Clean Water Act) In accordance with this law, process plants in the United States treat wastewater at on-site or near-site treatment centers before releasing it into the surrounding environment.

Wastewater effluent streams, along with water runoff from around the plant, are treated to control for pH, toxicity, suspended solids, and biological oxygen demand (for aquatic life protection) prior to discharge. Each of these controls is typically addressed with a separate method. Acidity and basicity is balanced through the addition of an acid or alkaline solution. Toxic wastewater may be treated with chemical processes or simply diluted to safe concentrations. Suspended solids can be removed via filtration and/or with clarifiers. Oxygen demand of wastewater can be mitigated using activated sludge treatment processes. Once the water quality complies with the EPA, and state-mandated, regulations, it can be safely released. More information on the large number of industry-specific guidelines for waste effluent can be found on the EPA website (http://www.epa.gov/eg/industrial-effluent-guidelines).

=Air-Pollution Management=
===Introduction===

In the United States air pollution is regulated in the Clean Air Act, and almost all pollutant emitting plants are regulated under this law. The types of plants that can release significant emissions include petroleum refineries, sulfur recovery plants, carbon-black plants, fuel conversion plants, chemical process plants, fossil fuel plants, and petroleum storage and transfer facilities. To receive permission to construct a plant must undergo a review to show that it will not cause a violation of the Ambient Air Quality Standards(Peters, 2003).

==Methods==

There are two major types of pollutants that are released into the air, particulates and and gaseous pollutants. Particulates can be removed with mechanical forces while gaseous pollutants typically need to removed by chemical or physical means (Peters, 2003).

Particulates and volatile pollutants that need to be removed before disposal may be present. Particle removal equipment includes: cyclones, wet scrubbers, electrostatic precipitators, and fabric-filter systems (Seider 2009) such as bag filters (. The two charts below are from ''Plant Design and Economics for Chemical Engineers'' and show the types of equipment, separation methods, and particle sizes in different pollutant separation technologies (Peters, 1991).

[[File:AirPolChart.png|thumb|border|center|middle|upright=4|link=|atl=|Different Methods of Particle Separation, the Particle Sizes they Can Remove, and the Technologies Used]]

[[File:AirPolTable.png|thumb|border|center|middle|upright=4|link=|atl=| This Chart Shows the Specifications and Limitations of Different Separations Technologies Including Particle Size, Efficiencies, and Temperatures(Peters, 1991)]]

Methods for removing inorganic and organic gaseous pollutants include: absorption, adsorption, condensation, and combustion (Seider 2009 pg 609). A list of typical of gases pollutants and their sources from ''Plant Design and Economics for Chemical Engineers'' is shown below (Peters, 1991).

[[File:GaseousPollutants.png|thumb|border|center|middle|upright=4|link=|atl=| Common Gaseous Pollutants and their Sources]]

Typically Gas-liquid absorption processes are done completed in a vertical, countercurrent, flow through packed, plate, or spray towers. These systems require good liquid-gas contact and proper equipment. These systems also often have significant energy consumption because of large pressure drops (Peters, 2003). For high volume systems absorption by scrubbing with water or another solvent is the most widely used method (Towler, 2012). Dry adsorbents can be used to remove the last races of gaseous pollutants. Adsorption typically requires blowers, condensers, separators, and controls. You also typically need two packed beds so that one can be used while the other is regenerated. Examples of adsorbents are molecular sieves and activated carbon. Incineration is typically used when there are gas streams that have no recovery value. This can be done with direct flame or catalytic oxidation. Catalytic oxidation usually has higher capital costs, but lower operating costs because it does not require fuel.
===Wet Scrubbers===
Wet scrubber use lime or limestone and water to remove SO2 and acid gases. The mixture can be injected into a scrubber or the gases can be bubbled through this mixture. This results in removal of 90-98% of SO2 and and acid gases (Clean Coal Technologies).

===Dry Scrubbers===
Dry scrubber blow powdered adsorbents into a vessel with gases and then after it has captured the SO2 and acid gases it is separated from gas using a fabric filter. These systems remove 90-93% of the contaminants (Clean Coal Technologies).

===Low-NOx Burners===
The purpose of low NOx burners is to decrease the amount of NOx created when the coal is burned. This is done by injecting coal and air in boilers. This can result in 40-50% NOx reduction. If air is injected into the area above the burner this can actually cause almost 70% NOx reduction (Clean Coal Technologies).

===Selective Non-Catalytic and Catalytic Reductions===
These systems inject ammonia into gases to remove NOx. The Catalytic reductions add a catalyst to the ammonia being injected to aid in the removal. Con-catalytic reductions result in about 35% removal, but adding a catalyst can increase that amount to about 90%. The catalytic-reduction can remove up to 80% of mercury as well (Clean Coal Technologies).

===Fabric Filters===
Also known as baghouses,these [[Separation processes#Filtration| filters]] remove particulates by passing air through filters. These can separate as much as 99.9% of particulate matter (Clean Coal Technologies).

===Electrostatic Precipitators===
Electrostatic Precicpitators remove particulate matter as gas passes through a device that has charged metal plates. The particles are then removed because of static electricity. These systems can remove between 99-99.9% of particulate matter (Clean Coal Technologies).

===Super-critical Boilers===
Super-Critical Boilers and Ultra-Supercritical Boilers operate at temperature and pressure higher than regular boilers. By operating at higher temperature these systems become more efficient. Super Critical Boilers typically have 10%-20% CO2 emissions than other similar sub-critical technologies. Ultra-Supercritical boilers can be as much as 30% more efficent than sub-critical technologies (Clean Coal Technologies).

===More Information===

For more information regarding some of these separations equipment see [[Solids-involved equipment]]

For more information Regarding Cyclones see [[Separation processes#Cyclones| Cyclones]], and for modeling cyclones see their[[Solids-involved equipment#HYSYS Simulation |HYSYS Simulation]]

You can find more information on [[Separation processes#Absorption| Absorption]] and [[Separation processes#Adsorption| Adsorption]] in [[Separation processes]]

===Outcomes===

====United States Common Pollutant Emissions====

The United States implemented the Clean Air Act in 1970 and since then emissions in the U.S. have been drastically reduced. As a matter of fact despite increases in the population in the last 35 years the amount of pollutants emitted have decreased by almost 70% (EPA). More importantly this demonstrates that reasonable efforts can be put towards environmental protection without causing too much harm to industry.

[[File:EPA.png|thumb|border|center|middle|upright=4|link=|atl=| GDP and Other Growth Factors Vs Common Pollutants Released]]

===China Clean Coal===

====Success Through 2005====

Coal is a very inexpensive and abundant source of energy and is abundant in China (Xu 2010). In China Coal the cause of 90% of SO2 emissions, 70% of dust emissions, and 67% of NOx emissions, and 70% of CO2 emissions. These numbers are staggering considering that Coal has never been more than 50 percent of China's Energy Supply (Xu 2010).

[[File:Coal.jpg|thumb|border|center|middle|upright=4|link=|atl=| Coal Consumption for Power and the Percentage of Coal Used (Xu 2010)]]

Despite the increasing coal consumption, high efficiency electric dust removal systems with efficiencies that are as high as 99.6% have greatly decreased soot emissions to 32% below 1980's levels as of 2005.

[[File:Soot.jpg|thumb|border|center|middle|upright=4|link=|atl=| SO2 and Soot Emissions from 1981 to 2005 (Xu 2010)]]

====Technologies and Life Cycle Comparison====
As of 2010 China was consuming 48.2% of coal globally. Four potential technologies integrated gasification combined cycle (IGCC), sub-critical coal power generation (Sub-C); super-critical coal power generation (Super-C) ultrasuper-critical coal power generation (USC). These technologies are compared mostly on the basis of net generating efficiency and efficiency. Net generating efficiency is the output of the plant divided by the total available energy in the fuel used (Liang, 2013).

where efficiency is define as: η=E/Etotal

and

Etotal=Emining+Etransportation+Egeneration

[[File:technologies.PNG|thumb|border|center|middle|upright=4|link=|atl=|A comparison of many of the parameters important to lfe-cycle analysis (Liang, 2013)]]

This demonstrates that USC has the best net generation efficiency, while also having the largest capacity for a single system.

[[File:Lifecycle2.PNG|thumb|border|center|middle|upright=4|link=|atl=| (Liang, 2013)]]

This breaks down the energy efficiency of each system and demonstrates that USC is the most efficient

[[File:Lifecycle.PNG|thumb|border|center|middle|upright=4|link=|atl=| (Liang, 2013)]]

IGCC and USC are lowest on different element, but since these all have different global warming potential it is difficult to tell which is the most efficient.

[[File:GWP.PNG|thumb|border|center|middle|upright=4|link=|atl=| Environemntal measures and their explanations (Liang, 2013)]]

This table breaks down the different measures of which systems are better

[[File:Measures.PNG|thumb|border|center|middle|upright=4|link=|atl=| Environemntal measures and their explanations (Liang, 2013)]]

And the table above defines all the relevant terms.

[[File:Price.PNG|thumb|border|center|middle|upright=4|link=|atl=| Capital cost in $/kW of clean coal power-generation technologies (Liang, 2013)]]

This demonstrates that the capital cost per unit of energy production is very high for IGCC, but shows that Super-C and USC are very competitive.

All of the data suggests that USC is the highest in both energy efficiency and net generating efficiency while competitive in price. The IGCC while potentially the best on emissions is too much more expensive to use on a wide scale (Liang, 2013).

=Solid Waste=
U.S. federal regulations require that solid waste be classified as hazardous or nonhazardous. Conditions for a classification of hazardous include: ignitability, corrosivity, reactivity, toxicity, or posing a substantial threat to the surrounding environment and its inhabitants. Hazardous waste must be treated on- or near-site before being removed in containers. Non-hazardous waste may be landfilled or incinerated in some cases. A typical estimate of costs for waste disposal is $0.03/lb for nonhazardous solids and $0.10/lb for hazardous solids. (Seider 2009 pg 609)

=References=

*Biegler LT, Grossmann IE, Westerberg AW. ''Systematic Methods of Chemical Process Design'' Prentice-Hall: Upper Saddle River, 1997.

*Broughton, J. ''Process Utility Systems: Introduction to Design, Operation and Maintenance'' Institution of Chemical Engineers: Rugby, Warwickshire, UK, 1994.

*Clean Coal Technologies. America's Coalition for Clean Coal Electricity. http://www.americaspower.org/clean-coal-technologies-1663/

*Duke Energy Company (2013). How Do Coal Fired Plants Work? Charlotte: Duke Energy.

*"Economizers." Energy Solutions Center Inc.: Boiler Consortium. http://www.cleanboiler.org/Eff_Improve/Efficiency/Economizers.asp. Accessed February 21, 2016.

*Lehman B, Worrell E. ''Electricity Production from Natural Gas Pressure Recovery Using Expansion Turbines.'' Lawrence Berkeley National Laboratory; 2001.

*"Overview of the Clean Air Act and Air Pollution." Environmental Protection Agency. November 17, 2015. http://www.epa.gov/clean-air-act-overview. Accessed February 5, 2016.

*"Oxygen Control." Energy Solutions Center Inc.: Boiler Consortium. http://www.cleanboiler.org/Eff_Improve/Efficiency/Oxygen_Control.asp. Accessed February 21, 2016.

*Peters, Max S.; Timmerhaus, Klaus D.; West, Ronald E. (2003). " Plant Design and Economics for Chemical Engineers." McGraw Hill Higher Education.

*Ruina Xu Energy (2010). "Clean coal technology development in China." Elsevier.

*Seider, Seader, Lewin, Widagdo. (2009). ''Plant Design and Economics for Chemical Engineers, 5th Edition.'' Hoboken: Wiley.

*Seider, Seader, Lewin. (2008). ''Product and Process Design Principles, 2nd Edition.'' Hoboken: Wiley.

*Summary of the Clean Water Act. United States EPA website. http://www.epa.gov/laws-regulations/summary-clean-water-act

*Towler, G.P. and Sinnot, R. (2012). ''Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design.'' Elsevier.

*Turton R.; Bailie, R.C.; Whiting, W.B.; Shaeiwitz J.A.; Bhattacharyya D. (2012). ''Analysis, Synthesis, and Design of Chemical Processes.'' Upper Saddle River: Prentice Hall.

*G.D. Ulrich, A Guide to Chemical Engineering Process Design and Economics, Wiley: New York, 1984.

*Xiaoye Liang, Zhihua Wang, Zhijun Zhou, , Zhenyu Huang, Junhu Zhou, and Kefa Cen. (2013). "Up-to-date life cycle assessment and comparison study of clean coal power generation technologies in China" Elsevier.

Utility systems

2016-02-22T05:13:34Z

Bts511: /* Technologies and Life Cycle Comparison */

Authors: David Chen, [2014] Joshua Lee, [2015] Brett Sleyster, [2016] and Tom Aunins [2016] 

Stewards: David Chen, Jian Gong, and Fengqi You

Date Presented: January 13, 2014 /Date Revised: January 14, 2014

 


__TOC__

=Introduction=
Many chemical processes do not take place at ambient temperature or pressures. In order to reach these non-ambient conditions, utilities will have to be used to raise or lower temperatures and compress gases. (Towler, Towler/UOP) Utilities often contribute 5 to 10% of the price of a product, and may come from public or private utility companies or on-site plants. For purchased utilities, costs depend on consumption, while for company-owned utilities, there will be both capital and operating costs. They include things such as steam for heating, electricity, cooling water, refrigeration, fuels such as natural gas, wastewater treatment, waste disposal, and landfill. Steam is often the largest utility cost. Cogeneration unit can supply electricity accompanied with different steam pressures. (Seider 2009)

=Electricity=

Electricity is used to power many different kinds of equipment. It has many advantages: it is efficient (> 90%), reliable, available in a wide range of power, shaft speeds, designs, lifetimes, convenience, costs, and maintenance. It is generally used up to 200 hp, and sometimes over 10,000 Hp. In chemical process plants, the electricity demand is generally determined by the work or energy required for compression, pumping, air cooling, lights, and many other items. This electricity often times is purchased from local electricity providers, but many plants generate their own electricity via sophisticated processes.

Electricity is rarely used as a primary heat utility in large-scale chemical plants for a variety of reasons. The main disadvantages of using electricity as a heat utility are as follows (Towler 2012):

:*Heat from electricity is two to three times more expensive than heat from fuels. This is attributed to the lack of efficiency when creating heat from electricity.

:*Electrical heating units are expensive, require high maintenance, and must comply with strict safety regulations.

:*Electrical heating units are unsafe compared to steam heating units. In steam systems, the physically steam controls the temperature, whereas in electrical heating units temperature is controlled by temperature controllers, which can fail or burn out.

The use of electricity carries with it some hazards depending on the environment. Extra care must be taken when using electrically-powered equipment in areas which may have combustible fluids, vapors, or dust, and where liquids may be present. (Seider 2009 pg 606)

==Conventional Power Station==

In general, most electricity is generated from a conventional coal-fired process, whether it be on-site or purchased from a provider. Coal-fired processes have been used to create electricity throughout history, and technological advances have increased its efficiency and use worldwide. In a coal-fired steam station—much like a nuclear station—water is turned into steam, which in turn drives turbine generators to produce electricity. There are several variations on how to create energy from coal, but here are the basics of how a coal-fired process works:

[[File:coalfired.png|thumb|border|center|middle|upright=4|link=|atl=|General Coal-Fired Process Diagram (Duke Energy Company)]]

*Heat is created:
::Before the coal is burned, it is pulverized to the fineness of talcum powder. It is then mixed with hot air and blown into the firebox of the boiler. Burning in suspension, the coal/air mixture provides the most complete combustion and maximum heat possible.

*Water turns to steam:
::Highly purified water, pumped through pipes inside the boiler, is turned into steam by the heat. The steam reaches temperatures of up to 1,000 degrees Fahrenheit and pressures up to 3,500 pounds per square inch, and is piped to the turbine.

*Steam turns the turbine:
::The enormous pressure of the steam pushing against a series of giant turbine blades turns the turbine shaft. The turbine shaft is connected to the shaft of the generator, where magnets spin within wire coils to produce electricity.

*Steam is converted back to water:
::After doing its work in the turbine, the steam is drawn into a condenser, a large chamber in the basement of the power plant. In this important step, millions of gallons of cool water from a nearby source (such as a river or lake) are pumped through a network of tubes running through the condenser. The cool water in the tubes converts the steam back into water that can be used over and over again in the plant.

*Repeat:
::The cooling water is returned to its source without any contamination, and the steam water is returned to the boiler to repeat the cycle.

===Advantages & Disadvantages of Coal-Fired Energy Production===

[[File:coalfiredtable.png|thumb|border|center|middle|upright=4|link=|atl=|Advantages and Disadvantages of Coal-Fired Processes (Seider 2008)]]

==Gas-Turbine Cogeneration Process==

When generating energy on-site, many plants use a gas-turbine cogeneration process. The thermal efficiency of a gas-turbine process is in the range of 70-80% while conventional power stations, such as coal-fired processes, have a 30-40% efficiency. The lower efficiency in more conventional power stations is attributed to wasted heat in the exhaust steam in the condenser. One example of a gas-turbine process is outlined in the following figure. Figure 3.1 is a gas-turbine cogeneration process with a heat recovery steam generator (waste-heat) boiler.

[[File:cogeneration.png|thumb|border|center|middle|upright=4|link=|atl=|Gas-Turbine Cogeneration with a heat recovery steam generator boiler (Towler 2012 Fig 3.1)]]

Overall, the process illustrated is not much different from a coal-fired process. The main differences are that the cogeneration process creates both electricity and a heat utility, and cogeneration processes use natural gas instead of coal. Many of the advantages and disadvantages are similar to those of the coal-fired process, but the cogeneration has a much higher efficiency, creates heat to be used in another process, and uses a more volatile and expensive fuel. The main advantage of cogeneration over coal-fired energy production is that heat is not wasted. In coal-fired processes, heat is released and wasted during electricity generation. Cogeneration captures some, if not all of the byproduct for heat, which is an extremely useful utility that will be discussed in subsequent sections. In summary, the cogeneration plant is superior to the coal-fired process because of its higher efficiency and ability to create a useful heat utility.

Obviously any engineer would design the cogeneration plant to meet at least the energy requirement necessary for plant operation, but cogeneration plants often times are designed to exceed the plant electricity requirement to drive another source of capital. Many describe this scenario as a "make or buy" scenario (Towler 2012). This scenario provides chemical producers leverage when negotiating contracts with outsourced electricity providers and this allows plants to purchase electricity at a wholesale price. This is a huge advantage of considering on-site electricity production because electricity is needed in relatively high quantities for all chemical plants. Being able to minimize electricity costs, or even profit off of electricity production is a huge economical consideration that all plants employ.

=Process Heating=
The key objective of this section is to discuss how processes are heated. Heating utilities are necessary for proper usage of distillers, reactors, condensers, and several other integral types of equipment. More specifically, steam, fired heat, and hot oil/specialized heat transfer fluids will be discussed in the following subsections.

==Steam==
Steam is the most commonly used heat utility used in chemical plants, and as a result understanding how it is used is essential in the study of Utility systems. Steam is used both as a process fluid (feedstock, diluent to absorb heat of reaction, heating agent, and stripping agent in absorbers and adsorbers ) and utility. It can be used to drive pumps and compressors, ejectors (for producing a vacuum), and heat exchangers. As one can clearly see, steam is a versatile, and useful utility.

Here are a few advantages of using steam as opposed to other methods of process heating (Towler 2012):

:*By controlling the pressure of the steam, one can control the temperature at which the heat is released. Having a strong control over the temperature is essential in several processes. =
:*Steam is an efficient heat source because the heat of condensation of steam is very high. Meaning that there is is high output per mass of utility at a constant temperature.
:*Heat exchangers that use steam are relatively cheap because condensing steam has a high heat transfer coefficient.
:*Steam is nonflammable, nontoxic, and inert to several process fluids.

Chemical plants generally have a network of pipelines exclusively for providing steam. These networks generally have steam at a low pressure, a medium pressure, and a high pressure. The image below illustrates a basic steam system.

[[File:Steam.png|thumb|border|center|middle|upright=4|link=|atl=|(Towler 2012 Fig 3.2)]]

In the diagram above, boiler feed water at a high pressure is preheated and fed to other boilers. These other boilers superheat the steam to create a high pressure and high temperature steam stream. The steam is superheated past the dew point to account for heat loss in the pipelines. A portion of the high pressure steam is used for process heating in areas of the plant that require high temperatures. The rest of the high pressure steam is turned into medium pressure steam by valves and steam turbines. The medium pressure steam is then used to heat medium temperature processes and to form low pressure steam. The low pressure steam can be used to heat low pressure processes and it can be expanded in condensing turbines to create shaft work and energy. In summary, steam can be used for an innumerable amount of action items in a plant. High pressure, medium pressure, and low pressure steam can all be used as a heat source. Low pressure steam has utilities in creating electricity and it also has several other uses.

==Fired Heat==

In many cases, processes in a plant require a heat source stronger than high pressure and temperature steam. That is when fired heat is used, which is generally at temperatures above 523K. Streams can be heated directly in the furnace tubes or via a hot oil circuit or heat transfer fluid, which will be discussed in detail in the next section. Most fired heaters use natural gas as fuel because it burns cleaner than fuel oil. A cleaner burning fuel is always advantageous due to environmental and safety concerns. Furthermore, natural gases usually result in less wear and tear in burners and fuel lines.

Depending on the application of the fired-heater, different design specifications can be implemented to make the fired-heater as efficient as possible. The basic construction of a fired heater starts with a cylindrical chamber that is lined with with refractory bricks. The stream that is to be heated flows through tubes inside of the furnace. These tubes can be arranged in several different arrangements such as, around the walls of the furnace, or in horizontal or vertical banks (Towler 2012). The figure below illustrates the basic construction of the fired-heater and varying tube arrangements.

[[File:firedheater.png|thumb|border|center|middle|upright=4|link=|atl=|Basic Fired-Heater Diagrams (Towler 2012 Figure 19.66)]]

Fuel is burned to heat the entire furnace, and the heat transfer occurs from the combustion gases inside of the furnace across the tubes that are filled with our desired stream. The heat transfer between the tubes and the furnace is accomplished primarily via radiation. Modern designs take advantage of convective heat transfer by adding a smaller section on top of the fired-heater where the combustion gases flow over banks of tubes as seen in (c) in the figure above. Heat transfer can be further improved via convection by adding fins or pins in the combustion section.

The cost of fired heat can be calculated by the cost of fuel fired. Natural gas and heating oil are traded as commodities and prices can be found at many online trading sites or business news sites (i.e., www.cnn.money.com). Past historic prices for forecasting can be found in the Oil and Gas Journal or from the U.S. Energy Information Adminstration (www.eia.gov).

===Fuels===
Fuel is burned in utility facilities such as boilers, electricity generation, and cogeneration, and can be in solid, liquid, or gas form. It can also be burned to provide heating for a process or stream or to drive pumps and compressors. The fuel is usually burned with excess air to ensure complete combustion.

A way of quantifying the amount of heat generated is by using the heating values. Higher heating value (HHV) and the lower heating value (LHV) are used. The heating is the total heat evolved by complete combustion of a fuel with dry air with both at 60 ⁰F and the flue gas after combustion brought back down to ⁰F. If the water vapor in the flue gas is not condensed, we obtain the LHV. If the water vapor is condensed, the value of heat evolved is a bit higher, and this is the HHV.
Heating values for solids and liquids are usually on a per-mass basis, and gases on a per-volume basis.
A specified amount of heating cannot be met with the amount of fuel calculated using only the HHV. There will be heat losses, the flue gas temperature will be greater than 60 ⁰F, and water in the flue gas will typically be vapor. (Seider 608)

==Hot Oil/Specialized Heat Transfer Fluids==

Specialized heat transfer fluids and hot oil circuits are used as heat sources when steam and fired heat is not appropriate. Specialized heat transfer fluids and hot oil circuits are extremely versatile in that they can be used in the temperature range of 323K to 673K. This range however is quite variable. For hot oils, the upper temperature limit is gauged based off of the thermal decomposition of the oil and coking/fouling of heat exchanger tubes.

Hot oil circuit systems are most commonly used when the plant has several small temperature heating requirements because it is more economically sound. Rather than having several fired heaters heat each small temperature requirement, it is much more economical to have one fired heater heat the hot oil and circulate that oil through each of the process to meet all of the heat needs. Hot oil systems are also generally favored over high pressure steam in processes that involve high pressure differentials between the process stream and high pressure steam. Hot oil systems are favored in this scenario because of safety concerns. If the steam were to leak, the pressure drop could cause serious safety issues.

Mineral oils are the most commonly used heat transfer fluids, and one prominent example is Dowtherm A. Dowtherm A is a combination of 26.5 wt% diphenyl in diphenyl oxide (Towler 2012) and is extremely thermally stable. These mineral oil systems generally require high flow rates.

=Process Cooling=

==Cooling Water==
Cooling water is used to cool and/or condense streams. Cooling water is usually circulated between process heat exchangers and a cooling tower. Water is cooled during downward motion by contact with air blown upwards, which can bring the water temperature to come within ~ 5 ⁰F of air’s wet-bulb temperature.Approximately 80% of the temperature reduction is due to evaporation of the cooling water and heat transfer to the surrounding air. Water can also be cooled in spray ponds and cooling ponds. Both work by providing high area for water to exchange heat with air.
Water in cooling towers is lost through drift and blowdown, and makeup is usually 1.5 to 3% of the circulating rate.
If a large natural body of water is nearby, it can be used as a source of cooling water and discharged downstream. This water is usually filtered to remove salts and impurities that may lead to fouling, but it is not treated.

==Refrigeration==
Cooling water can usually be used to cool a stream to ~ 100 ⁰F. Air can only cool to ~ 120 ⁰F. Air may be used in places where water is scarce or more costly to transport. To cool or condense streams below 100 ⁰F, a refrigerant is typically used. Chilled brine can also be used, but is less common.

Until 1995, CFC Freon R-12 (dichlorodifuloromethane) and HCFC Freon R-22 (chlorodifuloromethane) were commonly used refrigerants. However, the chlorine atom in the molecules caused the depletion of the ozone layer. The U.S. Clean Air Act Amendments of 1990 went into effect in 1995, and the production of these refrigerants has since stopped or been greatly reduced.

Cost estimates are based on ton-day of refrigeration, where a ton is the heat that needs to be removed to freeze 1 ton per day of water at 32 ⁰F. Substitutes have since been developed. R-134a is often used in place of R-12. According to the EPA, R-134a is not combustible at ambient conditions, and is essentially non-toxic under 400 ppm, and is not ozone-depleting. (Seider pg 607)

=Energy Efficiency=
One of the chief concerns in selecting and designing process utility systems for heating and cooling is how to achieve the most energy efficient design. There are countless means by which plants lose energy, two of the foremost being through the mixing of different temperature or pressure streams and through the disposal of warmed cooling water. (Seider, 2009) Proper utilities design can help mitigate each of these losses as well as many others. The energy efficiency of a plant will depend primarily on the heating and cooling methods that are being used and the overall system design itself. These two parameters are important in determining how well energy is transferred to the desired media as well as how well that energy is recovered and recycled.

==Hot Utility Efficiency==
As mentioned above, the most commonly used utilities for process heating in large scale processes are steam, fired heat, and hot oil heaters. Of these, steam is the most commonly used. Electricity, while efficient at creating power, is not a viable source of heat in large industrial processes. Common ranges of heating efficiency for these three methods are displayed in Table 1. (Towler and Sinnott, 2012; Broughton, 1994)

{| class="wikitable"
|+Table 1: Process Heating Efficiencies
|-
! Process Heating Method
! Typical Efficiency
|-
| Steam (via coal boiler)
| 72%
|-
| Steam (via gas boiler)
| 73%
|-
| Steam (via oil boiler)
| 75%
|-
| Fired Heat w/ Convective Section
| 85%
|-
| Fired Heat w/o Convective Section
| 60%
|-
| Hot Oil Heaters/Vaporizers
| 80-85%
|}

===Steam===
As steam is so popular for heating purposes, it is useful to analyze the numerous ways in which losses can occur in steam systems. There are five primary sources of inefficiency and heat loss in the generation and distribution of steam throughout a process plant:

*The heat content of boiler exhaust gas
*Incomplete combustion of boiler fuel
*Radiant losses from the boiler exterior
*Blowdown losses
*Distribution losses (pipe transport, steam traps, etc.)

The first four of these losses take place at the boiler and contribute to the heating efficiencies seen in Table 1 for steam created with coal, gas, and oil. Several methods can be used to minimize these losses, one of the foremost being the control of air-to-fuel ratio in the boiler. This ratio must be managed by weighing losses due to uncombusted fuel against losses due to the heat content of excess exhaust gas. (Broughton, 1994)

[[File:Air-fuel ratio.PNG|thumb|border|center|middle|upright=4|link=|atl=|Air-to-fuel ratio relationship with heat losses (Broughton Fig 2.3)]]

The air-to-fuel ratio can be optimized using a feedback process controller. The control system will analyze the oxygen content of exhaust air and adjust the incoming air flow rate to achieve a set percentage of excess air. While desired excess oxygen will vary depending on the type of fuel, it is consistently seen that in the minimum loss-range a 1% increase in excess air will result in a 1% decrease in efficiency. ("Oxygen Control")

In addition to air-to-fuel ratio management, steam heat losses at the boiler can be mitigated via energy recovery, which is discussed in further detail [[Utility_systems#Energy_Recovery|below]].

Outside the boiler, losses in distribution of steam throughout facilities can decrease energy efficiency by a significant amount. This can cause up to a 60% increase in losses, but typically results in an overall steam efficiency of 50-55% (down from approximately 75% at the boiler). (Broughton, 1994) There are several ways that this issue can be addressed. First and foremost are steam metering systems, which can be used to monitor heat supply to separate sections of the process facility, check efficiency of fuel use, and determine which processes in a given plant have the highest steam requirements. Another method of minimizing transport losses is to decentralize steam generation systems. It can be advantageous to have numerous smaller boilers rather than a single centralized boiler so that steam does not have to travel as long of distances to reach its destination.

==Cold Utility Efficiency==
Efficiency in cooling processes depends more on the method used, and by extension the amount of coolant needed. Water and air utility efficiencies depend primarily on the fluid flow required to maintain the system at a desired temperature, while powered refrigeration utilities (for colder processes) have efficiencies at approximately 60%--but ranging up to 90%--of Carnot cycle efficiency, a metric of ideal refrigeration efficiency. (Towler and Sinnott, 2012) Cooling systems represent by definition a loss of energy from the main process to the utility stream, and as such it is often useful to find other uses for the heated media before discharge.

==Energy Recovery==
Recovery and recycle of energy is perhaps the most important aspect of creating an energy efficient plant design, and it is important for process engineers to fully consider possibilities for heat recovery in order to aid in economic viability.

===Process Heat Exchange===
Heat exchanger networks are a very common energy recovery method in industrial processes. These networks most frequently allow energy from heated product streams to be transferred to feed streams that must be brought up to process temperature. (Biegler, 1997) More information on the function and design of heat exchanger networks can be found on the [[Heat_exchanger#Heat_Exchanger_Networks|heat exchanger wiki page]]. The following are several examples of energy recovery via heat exchange that are used in industrial processes.

In distillation columns the bottoms and distillate effluents have the potential for energy exchange. Though the condenser at the top of the column cannot supply its waste heat to the reboiler due to their respective temperatures, the effluent streams can supply heat to the feed via a feed-effluent exchanger. This reduces the utility requirements to raise the feed to column temperature. (Biegler, 1997)

[[File:Feed_sterilization.jpg|thumb|border|right|middle|upright=4|link=|atl=|Feed sterilization schematic. (Towler and Sinnott Fig. 3.30)]]

Feed sterilization, commonly used in the food industry, is a common application for heat recovery through process stream heat exchange. In this application, the feed must be heated for a certain amount of time to kill any biological contaminants, after which it can be used to heat the new raw feed for sterilization. This reduces energy demands on the steam heater and thus reduces cost. (Towler and Sinnott, 2012)

In multi-vessel batch processes it can be advantageous to exchange heat as the process fluid is being transferred between vessels. Like the previous examples, this reduces the utility needed to bring the colder feed up to process temperature, thus reducing costs. (Towler and Sinnott, 2012)

[[File:Batch_heat_exchange.jpg|thumb|border|center|middle|upright=4|link=|atl=|Multi-vessel batch heat exchange schematic (Towler and Sinnott Fig. 3.31)]]

===Utility Regeneration===
[[File:Waste_heat_boiler.jpg|thumb|border|right|middle|200px|upright=4|link=|atl=|Industrial modular HRSG]]
When recovery of waste heat via transfer to other process streams is inconvenient or impossible, energy efficiency can still be improved through the regeneration of utilities. This is commonly done through the regeneration of steam by removing heat from exiting streams or from highly exothermic reactions. Waste heat in exiting streams can be removed via heat recovery steam generators (HRSGs), and is most often used on exiting gas streams. Heat recovery from reactions is a viable option when the reactor temperature will be at 150 C or above, as this will create steam at high enough pressure to be used in other processes. (Towler and Sinnott, 2012)

In the case of steam, waste heat and water treatment losses can be recovered from the utility generation process itself. One of the most common ways to do this is with an economizer. As seen in the schematic below, economizers have heat exchangers fit to the exhaust gas flow in order to transfer waste heat from these gases to the incoming boiler feed water. This can result in fuel energy savings of approximately 15% for typical excess air percentages in steam boilers. (Broughton, 1994; "Economizers")

[[File:Economizer.png|thumb|border|center|middle|upright=4|link=|atl=|Economizer schematic ("Economizers") and economizer energy recovery correlations (Broughton Fig 2.2)]]

[[File:Condensate recycle.PNG|thumb|border|right|middle|200px|upright=4|link=|atl=|Condensate return fuel savings (Broughton Fig 2.4)]]

Additionally, there are still more opportunities for recovery in steam generation processes. One typical case is the use of waste heat to preheat air entering the boiler. Optimistically, this can result in a 5% improvement in heat recovery for the system. Another method is the recycle of steam condensate back to the boiler. This has two main benefits: it retains the heat still present in the condensed steam and it retains the pretreatment chemicals added to the steam. (Broughton, 1994)

Furthermore, there is opportunity for energy recovery in the expansion of compressed gas through a turbine to create electricity, a process that can be economically viable given sufficiently high flows or pressure. One of the primary application of this type of energy recovery is in the creation of medium pressure and low pressure steam. In most processes, all steam is generated as high pressure steam and can be expanded through a turbine to decrease its pressure. Such technology has also been used in processes to synthesize ammonia, perform air separations, and synthesize nitric acid. (Towler and Sinnott, 2012) Recently, however, there has been a particularly strong interest for energy recovery in the natural gas industry, when gas is decompressed from major pipelines to residential low-pressure pipelines. A 2001 study estimated that there is the potential to recover 21 TWh, representing 11% of natural gas transport energy, via gas expansion. (Lehman)

=Process water and boiler-feed water=
Process water is water that will be directly used in the process. Boiler-feed water (BFW) is used to produce steam. Both may need to be purified to prevent impurities from contaminating a process or from foul equipment. It can be used as a cooling stream when the temperature of the stream to be cooled is greater than ~300 ⁰F. Cost of BFW can be partially offset by the steam credit.

Process water that undergoes moderate pretreatment can cost ~ $0.75/1,000 gal.

Extensive treatment ~ $6.00/1,000 gal.

Sterilized for pharmaceutical processes ~ $550/1,000 gal. (Seider pg 608)

=Demineralized Water=
In demineralized water, minerals have been removed by ion exchange. In boiler feed water, this is to prevent salt deposition, corrosion, formation of foam, and sluicing. In process water, the ions may contaminate the process.

=Waste Treatment=
Most chemical processes will produce some sort of waste. Disposal occurs to the atmosphere (in the case of some gases), sewers, body of water, or a landfill. Waste may require some treatment before disposal to meet regulations. Depending on process economics, byproducts may be recovered and processed. (Seider 2009 pg 609)

==Wastewater Treatment==
The United States EPA regulates industrial wastewater disposal through the Clean Water Act, introduced in 1948 as the Federal Water Pollution Control Act and amended to its current form in 1972. The sweeping 1972 amendments allowed the EPA to prevent industries and persons from discharging contaminated water into fresh water sources and set water quality standards. (Summary of the Clean Water Act) In accordance with this law, process plants in the United States treat wastewater at on-site or near-site treatment centers before releasing it into the surrounding environment.

Wastewater effluent streams, along with water runoff from around the plant, are treated to control for pH, toxicity, suspended solids, and biological oxygen demand (for aquatic life protection) prior to discharge. Each of these controls is typically addressed with a separate method. Acidity and basicity is balanced through the addition of an acid or alkaline solution. Toxic wastewater may be treated with chemical processes or simply diluted to safe concentrations. Suspended solids can be removed via filtration and/or with clarifiers. Oxygen demand of wastewater can be mitigated using activated sludge treatment processes. Once the water quality complies with the EPA, and state-mandated, regulations, it can be safely released. More information on the large number of industry-specific guidelines for waste effluent can be found on the EPA website (http://www.epa.gov/eg/industrial-effluent-guidelines).

=Air-Pollution Management=
===Introduction===

In the United States air pollution is regulated in the Clean Air Act, and almost all pollutant emitting plants are regulated under this law. The types of plants that can release significant emissions include petroleum refineries, sulfur recovery plants, carbon-black plants, fuel conversion plants, chemical process plants, fossil fuel plants, and petroleum storage and transfer facilities. To receive permission to construct a plant must undergo a review to show that it will not cause a violation of the Ambient Air Quality Standards(Peters, 2003).

==Methods==

There are two major types of pollutants that are released into the air, particulates and and gaseous pollutants. Particulates can be removed with mechanical forces while gaseous pollutants typically need to removed by chemical or physical means (Peters, 2003).

Particulates and volatile pollutants that need to be removed before disposal may be present. Particle removal equipment includes: cyclones, wet scrubbers, electrostatic precipitators, and fabric-filter systems (Seider 2009) such as bag filters (. The two charts below are from ''Plant Design and Economics for Chemical Engineers'' and show the types of equipment, separation methods, and particle sizes in different pollutant separation technologies (Peters, 1991).

[[File:AirPolChart.png|thumb|border|center|middle|upright=4|link=|atl=|Different Methods of Particle Separation, the Particle Sizes they Can Remove, and the Technologies Used]]

[[File:AirPolTable.png|thumb|border|center|middle|upright=4|link=|atl=| This Chart Shows the Specifications and Limitations of Different Separations Technologies Including Particle Size, Efficiencies, and Temperatures(Peters, 1991)]]

Methods for removing inorganic and organic gaseous pollutants include: absorption, adsorption, condensation, and combustion (Seider 2009 pg 609). A list of typical of gases pollutants and their sources from ''Plant Design and Economics for Chemical Engineers'' is shown below (Peters, 1991).

[[File:GaseousPollutants.png|thumb|border|center|middle|upright=4|link=|atl=| Common Gaseous Pollutants and their Sources]]

Typically Gas-liquid absorption processes are done completed in a vertical, countercurrent, flow through packed, plate, or spray towers. These systems require good liquid-gas contact and proper equipment. These systems also often have significant energy consumption because of large pressure drops (Peters, 2003). For high volume systems absorption by scrubbing with water or another solvent is the most widely used method (Towler, 2012). Dry adsorbents can be used to remove the last races of gaseous pollutants. Adsorption typically requires blowers, condensers, separators, and controls. You also typically need two packed beds so that one can be used while the other is regenerated. Examples of adsorbents are molecular sieves and activated carbon. Incineration is typically used when there are gas streams that have no recovery value. This can be done with direct flame or catalytic oxidation. Catalytic oxidation usually has higher capital costs, but lower operating costs because it does not require fuel.
===Wet Scrubbers===
Wet scrubber use lime or limestone and water to remove SO2 and acid gases. The mixture can be injected into a scrubber or the gases can be bubbled through this mixture. This results in removal of 90-98% of SO2 and and acid gases (Clean Coal Technologies).

===Dry Scrubbers===
Dry scrubber blow powdered adsorbents into a vessel with gases and then after it has captured the SO2 and acid gases it is separated from gas using a fabric filter. These systems remove 90-93% of the contaminants (Clean Coal Technologies).

===Low-NOx Burners===
The purpose of low NOx burners is to decrease the amount of NOx created when the coal is burned. This is done by injecting coal and air in boilers. This can result in 40-50% NOx reduction. If air is injected into the area above the burner this can actually cause almost 70% NOx reduction (Clean Coal Technologies).

===Selective Non-Catalytic and Catalytic Reductions===
These systems inject ammonia into gases to remove NOx. The Catalytic reductions add a catalyst to the ammonia being injected to aid in the removal. Con-catalytic reductions result in about 35% removal, but adding a catalyst can increase that amount to about 90%. The catalytic-reduction can remove up to 80% of mercury as well (Clean Coal Technologies).

===Fabric Filters===
Also known as baghouses,these [[Separation processes#Filtration| filters]] remove particulates by passing air through filters. These can separate as much as 99.9% of particulate matter (Clean Coal Technologies).

===Electrostatic Precipitators===
Electrostatic Precicpitators remove particulate matter as gas passes through a device that has charged metal plates. The particles are then removed because of static electricity. These systems can remove between 99-99.9% of particulate matter (Clean Coal Technologies).

===Super-critical Boilers===
Super-Critical Boilers and Ultra-Supercritical Boilers operate at temperature and pressure higher than regular boilers. By operating at higher temperature these systems become more efficient. Super Critical Boilers typically have 10%-20% CO2 emissions than other similar sub-critical technologies. Ultra-Supercritical boilers can be as much as 30% more efficent than sub-critical technologies (Clean Coal Technologies).

===More Information===

For more information regarding some of these separations equipment see [[Solids-involved equipment]]

For more information Regarding Cyclones see [[Separation processes#Cyclones| Cyclones]], and for modeling cyclones see their[[Solids-involved equipment#HYSYS Simulation|HYSYS Simulation]]

You can find more information on [[Separation processes#Absorption| Absorption]] and [[Separation processes#Adsorption| Adsorption]] in [[Separation processes]]

===Outcomes===

====United States Common Pollutant Emissions====

The United States implemented the Clean Air Act in 1970 and since then emissions in the U.S. have been drastically reduced. As a matter of fact despite increases in the population in the last 35 years the amount of pollutants emitted have decreased by almost 70% (EPA). More importantly this demonstrates that reasonable efforts can be put towards environmental protection without causing too much harm to industry.

[[File:EPA.png|thumb|border|center|middle|upright=4|link=|atl=| GDP and Other Growth Factors Vs Common Pollutants Released]]

===China Clean Coal===

====Success Through 2005====

Coal is a very inexpensive and abundant source of energy and is abundant in China (Xu 2010). In China Coal the cause of 90% of SO2 emissions, 70% of dust emissions, and 67% of NOx emissions, and 70% of CO2 emissions. These numbers are staggering considering that Coal has never been more than 50 percent of China's Energy Supply (Xu 2010).

[[File:Coal.jpg|thumb|border|center|middle|upright=4|link=|atl=| Coal Consumption for Power and the Percentage of Coal Used (Xu 2010)]]

Despite the increasing coal consumption, high efficiency electric dust removal systems with efficiencies that are as high as 99.6% have greatly decreased soot emissions to 32% below 1980's levels as of 2005.

[[File:Soot.jpg|thumb|border|center|middle|upright=4|link=|atl=| SO2 and Soot Emissions from 1981 to 2005 (Xu 2010)]]

====Technologies and Life Cycle Comparison====
As of 2010 China was consuming 48.2% of coal globally. Four potential technologies integrated gasification combined cycle (IGCC), sub-critical coal power generation (Sub-C); super-critical coal power generation (Super-C) ultrasuper-critical coal power generation (USC). These technologies are compared mostly on the basis of net generating efficiency and efficiency. Net generating efficiency is the output of the plant divided by the total available energy in the fuel used (Liang, 2013).

where efficiency is define as: η=E/Etotal

and

Etotal=Emining+Etransportation+Egeneration

[[File:technologies.PNG|thumb|border|center|middle|upright=4|link=|atl=|A comparison of many of the parameters important to lfe-cycle analysis (Liang, 2013)]]

This demonstrates that USC has the best net generation efficiency, while also having the largest capacity for a single system.

[[File:Lifecycle2.PNG|thumb|border|center|middle|upright=4|link=|atl=| (Liang, 2013)]]

This breaks down the energy efficiency of each system and demonstrates that USC is the most efficient

[[File:Lifecycle.PNG|thumb|border|center|middle|upright=4|link=|atl=| (Liang, 2013)]]

IGCC and USC are lowest on different element, but since these all have different global warming potential it is difficult to tell which is the most efficient.

[[File:GWP.PNG|thumb|border|center|middle|upright=4|link=|atl=| Environemntal measures and their explanations (Liang, 2013)]]

This table breaks down the different measures of which systems are better

[[File:Measures.PNG|thumb|border|center|middle|upright=4|link=|atl=| Environemntal measures and their explanations (Liang, 2013)]]

And the table above defines all the relevant terms.

[[File:Price.PNG|thumb|border|center|middle|upright=4|link=|atl=| Capital cost in $/kW of clean coal power-generation technologies (Liang, 2013)]]

This demonstrates that the capital cost per unit of energy production is very high for IGCC, but shows that Super-C and USC are very competitive.

All of the data suggests that USC is the highest in both energy efficiency and net generating efficiency while competitive in price. The IGCC while potentially the best on emissions is too much more expensive to use on a wide scale (Liang, 2013).

=Solid Waste=
U.S. federal regulations require that solid waste be classified as hazardous or nonhazardous. Conditions for a classification of hazardous include: ignitability, corrosivity, reactivity, toxicity, or posing a substantial threat to the surrounding environment and its inhabitants. Hazardous waste must be treated on- or near-site before being removed in containers. Non-hazardous waste may be landfilled or incinerated in some cases. A typical estimate of costs for waste disposal is $0.03/lb for nonhazardous solids and $0.10/lb for hazardous solids. (Seider 2009 pg 609)

=References=

*Biegler LT, Grossmann IE, Westerberg AW. ''Systematic Methods of Chemical Process Design'' Prentice-Hall: Upper Saddle River, 1997.

*Broughton, J. ''Process Utility Systems: Introduction to Design, Operation and Maintenance'' Institution of Chemical Engineers: Rugby, Warwickshire, UK, 1994.

*Clean Coal Technologies. America's Coalition for Clean Coal Electricity. http://www.americaspower.org/clean-coal-technologies-1663/

*Duke Energy Company (2013). How Do Coal Fired Plants Work? Charlotte: Duke Energy.

*"Economizers." Energy Solutions Center Inc.: Boiler Consortium. http://www.cleanboiler.org/Eff_Improve/Efficiency/Economizers.asp. Accessed February 21, 2016.

*Lehman B, Worrell E. ''Electricity Production from Natural Gas Pressure Recovery Using Expansion Turbines.'' Lawrence Berkeley National Laboratory; 2001.

*"Overview of the Clean Air Act and Air Pollution." Environmental Protection Agency. November 17, 2015. http://www.epa.gov/clean-air-act-overview. Accessed February 5, 2016.

*"Oxygen Control." Energy Solutions Center Inc.: Boiler Consortium. http://www.cleanboiler.org/Eff_Improve/Efficiency/Oxygen_Control.asp. Accessed February 21, 2016.

*Peters, Max S.; Timmerhaus, Klaus D.; West, Ronald E. (2003). " Plant Design and Economics for Chemical Engineers." McGraw Hill Higher Education.

*Ruina Xu Energy (2010). "Clean coal technology development in China." Elsevier.

*Seider, Seader, Lewin, Widagdo. (2009). ''Plant Design and Economics for Chemical Engineers, 5th Edition.'' Hoboken: Wiley.

*Seider, Seader, Lewin. (2008). ''Product and Process Design Principles, 2nd Edition.'' Hoboken: Wiley.

*Summary of the Clean Water Act. United States EPA website. http://www.epa.gov/laws-regulations/summary-clean-water-act

*Towler, G.P. and Sinnot, R. (2012). ''Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design.'' Elsevier.

*Turton R.; Bailie, R.C.; Whiting, W.B.; Shaeiwitz J.A.; Bhattacharyya D. (2012). ''Analysis, Synthesis, and Design of Chemical Processes.'' Upper Saddle River: Prentice Hall.

*G.D. Ulrich, A Guide to Chemical Engineering Process Design and Economics, Wiley: New York, 1984.

*Xiaoye Liang, Zhihua Wang, Zhijun Zhou, , Zhenyu Huang, Junhu Zhou, and Kefa Cen. (2013). "Up-to-date life cycle assessment and comparison study of clean coal power generation technologies in China" Elsevier.

Utility systems

2016-02-22T05:11:00Z

Bts511: /* Technologies and Life Cycle Comparison */

Authors: David Chen, [2014] Joshua Lee, [2015] Brett Sleyster, [2016] and Tom Aunins [2016] 

Stewards: David Chen, Jian Gong, and Fengqi You

Date Presented: January 13, 2014 /Date Revised: January 14, 2014

 


__TOC__

=Introduction=
Many chemical processes do not take place at ambient temperature or pressures. In order to reach these non-ambient conditions, utilities will have to be used to raise or lower temperatures and compress gases. (Towler, Towler/UOP) Utilities often contribute 5 to 10% of the price of a product, and may come from public or private utility companies or on-site plants. For purchased utilities, costs depend on consumption, while for company-owned utilities, there will be both capital and operating costs. They include things such as steam for heating, electricity, cooling water, refrigeration, fuels such as natural gas, wastewater treatment, waste disposal, and landfill. Steam is often the largest utility cost. Cogeneration unit can supply electricity accompanied with different steam pressures. (Seider 2009)

=Electricity=

Electricity is used to power many different kinds of equipment. It has many advantages: it is efficient (> 90%), reliable, available in a wide range of power, shaft speeds, designs, lifetimes, convenience, costs, and maintenance. It is generally used up to 200 hp, and sometimes over 10,000 Hp. In chemical process plants, the electricity demand is generally determined by the work or energy required for compression, pumping, air cooling, lights, and many other items. This electricity often times is purchased from local electricity providers, but many plants generate their own electricity via sophisticated processes.

Electricity is rarely used as a primary heat utility in large-scale chemical plants for a variety of reasons. The main disadvantages of using electricity as a heat utility are as follows (Towler 2012):

:*Heat from electricity is two to three times more expensive than heat from fuels. This is attributed to the lack of efficiency when creating heat from electricity.

:*Electrical heating units are expensive, require high maintenance, and must comply with strict safety regulations.

:*Electrical heating units are unsafe compared to steam heating units. In steam systems, the physically steam controls the temperature, whereas in electrical heating units temperature is controlled by temperature controllers, which can fail or burn out.

The use of electricity carries with it some hazards depending on the environment. Extra care must be taken when using electrically-powered equipment in areas which may have combustible fluids, vapors, or dust, and where liquids may be present. (Seider 2009 pg 606)

==Conventional Power Station==

In general, most electricity is generated from a conventional coal-fired process, whether it be on-site or purchased from a provider. Coal-fired processes have been used to create electricity throughout history, and technological advances have increased its efficiency and use worldwide. In a coal-fired steam station—much like a nuclear station—water is turned into steam, which in turn drives turbine generators to produce electricity. There are several variations on how to create energy from coal, but here are the basics of how a coal-fired process works:

[[File:coalfired.png|thumb|border|center|middle|upright=4|link=|atl=|General Coal-Fired Process Diagram (Duke Energy Company)]]

*Heat is created:
::Before the coal is burned, it is pulverized to the fineness of talcum powder. It is then mixed with hot air and blown into the firebox of the boiler. Burning in suspension, the coal/air mixture provides the most complete combustion and maximum heat possible.

*Water turns to steam:
::Highly purified water, pumped through pipes inside the boiler, is turned into steam by the heat. The steam reaches temperatures of up to 1,000 degrees Fahrenheit and pressures up to 3,500 pounds per square inch, and is piped to the turbine.

*Steam turns the turbine:
::The enormous pressure of the steam pushing against a series of giant turbine blades turns the turbine shaft. The turbine shaft is connected to the shaft of the generator, where magnets spin within wire coils to produce electricity.

*Steam is converted back to water:
::After doing its work in the turbine, the steam is drawn into a condenser, a large chamber in the basement of the power plant. In this important step, millions of gallons of cool water from a nearby source (such as a river or lake) are pumped through a network of tubes running through the condenser. The cool water in the tubes converts the steam back into water that can be used over and over again in the plant.

*Repeat:
::The cooling water is returned to its source without any contamination, and the steam water is returned to the boiler to repeat the cycle.

===Advantages & Disadvantages of Coal-Fired Energy Production===

[[File:coalfiredtable.png|thumb|border|center|middle|upright=4|link=|atl=|Advantages and Disadvantages of Coal-Fired Processes (Seider 2008)]]

==Gas-Turbine Cogeneration Process==

When generating energy on-site, many plants use a gas-turbine cogeneration process. The thermal efficiency of a gas-turbine process is in the range of 70-80% while conventional power stations, such as coal-fired processes, have a 30-40% efficiency. The lower efficiency in more conventional power stations is attributed to wasted heat in the exhaust steam in the condenser. One example of a gas-turbine process is outlined in the following figure. Figure 3.1 is a gas-turbine cogeneration process with a heat recovery steam generator (waste-heat) boiler.

[[File:cogeneration.png|thumb|border|center|middle|upright=4|link=|atl=|Gas-Turbine Cogeneration with a heat recovery steam generator boiler (Towler 2012 Fig 3.1)]]

Overall, the process illustrated is not much different from a coal-fired process. The main differences are that the cogeneration process creates both electricity and a heat utility, and cogeneration processes use natural gas instead of coal. Many of the advantages and disadvantages are similar to those of the coal-fired process, but the cogeneration has a much higher efficiency, creates heat to be used in another process, and uses a more volatile and expensive fuel. The main advantage of cogeneration over coal-fired energy production is that heat is not wasted. In coal-fired processes, heat is released and wasted during electricity generation. Cogeneration captures some, if not all of the byproduct for heat, which is an extremely useful utility that will be discussed in subsequent sections. In summary, the cogeneration plant is superior to the coal-fired process because of its higher efficiency and ability to create a useful heat utility.

Obviously any engineer would design the cogeneration plant to meet at least the energy requirement necessary for plant operation, but cogeneration plants often times are designed to exceed the plant electricity requirement to drive another source of capital. Many describe this scenario as a "make or buy" scenario (Towler 2012). This scenario provides chemical producers leverage when negotiating contracts with outsourced electricity providers and this allows plants to purchase electricity at a wholesale price. This is a huge advantage of considering on-site electricity production because electricity is needed in relatively high quantities for all chemical plants. Being able to minimize electricity costs, or even profit off of electricity production is a huge economical consideration that all plants employ.

=Process Heating=
The key objective of this section is to discuss how processes are heated. Heating utilities are necessary for proper usage of distillers, reactors, condensers, and several other integral types of equipment. More specifically, steam, fired heat, and hot oil/specialized heat transfer fluids will be discussed in the following subsections.

==Steam==
Steam is the most commonly used heat utility used in chemical plants, and as a result understanding how it is used is essential in the study of Utility systems. Steam is used both as a process fluid (feedstock, diluent to absorb heat of reaction, heating agent, and stripping agent in absorbers and adsorbers ) and utility. It can be used to drive pumps and compressors, ejectors (for producing a vacuum), and heat exchangers. As one can clearly see, steam is a versatile, and useful utility.

Here are a few advantages of using steam as opposed to other methods of process heating (Towler 2012):

:*By controlling the pressure of the steam, one can control the temperature at which the heat is released. Having a strong control over the temperature is essential in several processes. =
:*Steam is an efficient heat source because the heat of condensation of steam is very high. Meaning that there is is high output per mass of utility at a constant temperature.
:*Heat exchangers that use steam are relatively cheap because condensing steam has a high heat transfer coefficient.
:*Steam is nonflammable, nontoxic, and inert to several process fluids.

Chemical plants generally have a network of pipelines exclusively for providing steam. These networks generally have steam at a low pressure, a medium pressure, and a high pressure. The image below illustrates a basic steam system.

[[File:Steam.png|thumb|border|center|middle|upright=4|link=|atl=|(Towler 2012 Fig 3.2)]]

In the diagram above, boiler feed water at a high pressure is preheated and fed to other boilers. These other boilers superheat the steam to create a high pressure and high temperature steam stream. The steam is superheated past the dew point to account for heat loss in the pipelines. A portion of the high pressure steam is used for process heating in areas of the plant that require high temperatures. The rest of the high pressure steam is turned into medium pressure steam by valves and steam turbines. The medium pressure steam is then used to heat medium temperature processes and to form low pressure steam. The low pressure steam can be used to heat low pressure processes and it can be expanded in condensing turbines to create shaft work and energy. In summary, steam can be used for an innumerable amount of action items in a plant. High pressure, medium pressure, and low pressure steam can all be used as a heat source. Low pressure steam has utilities in creating electricity and it also has several other uses.

==Fired Heat==

In many cases, processes in a plant require a heat source stronger than high pressure and temperature steam. That is when fired heat is used, which is generally at temperatures above 523K. Streams can be heated directly in the furnace tubes or via a hot oil circuit or heat transfer fluid, which will be discussed in detail in the next section. Most fired heaters use natural gas as fuel because it burns cleaner than fuel oil. A cleaner burning fuel is always advantageous due to environmental and safety concerns. Furthermore, natural gases usually result in less wear and tear in burners and fuel lines.

Depending on the application of the fired-heater, different design specifications can be implemented to make the fired-heater as efficient as possible. The basic construction of a fired heater starts with a cylindrical chamber that is lined with with refractory bricks. The stream that is to be heated flows through tubes inside of the furnace. These tubes can be arranged in several different arrangements such as, around the walls of the furnace, or in horizontal or vertical banks (Towler 2012). The figure below illustrates the basic construction of the fired-heater and varying tube arrangements.

[[File:firedheater.png|thumb|border|center|middle|upright=4|link=|atl=|Basic Fired-Heater Diagrams (Towler 2012 Figure 19.66)]]

Fuel is burned to heat the entire furnace, and the heat transfer occurs from the combustion gases inside of the furnace across the tubes that are filled with our desired stream. The heat transfer between the tubes and the furnace is accomplished primarily via radiation. Modern designs take advantage of convective heat transfer by adding a smaller section on top of the fired-heater where the combustion gases flow over banks of tubes as seen in (c) in the figure above. Heat transfer can be further improved via convection by adding fins or pins in the combustion section.

The cost of fired heat can be calculated by the cost of fuel fired. Natural gas and heating oil are traded as commodities and prices can be found at many online trading sites or business news sites (i.e., www.cnn.money.com). Past historic prices for forecasting can be found in the Oil and Gas Journal or from the U.S. Energy Information Adminstration (www.eia.gov).

===Fuels===
Fuel is burned in utility facilities such as boilers, electricity generation, and cogeneration, and can be in solid, liquid, or gas form. It can also be burned to provide heating for a process or stream or to drive pumps and compressors. The fuel is usually burned with excess air to ensure complete combustion.

A way of quantifying the amount of heat generated is by using the heating values. Higher heating value (HHV) and the lower heating value (LHV) are used. The heating is the total heat evolved by complete combustion of a fuel with dry air with both at 60 ⁰F and the flue gas after combustion brought back down to ⁰F. If the water vapor in the flue gas is not condensed, we obtain the LHV. If the water vapor is condensed, the value of heat evolved is a bit higher, and this is the HHV.
Heating values for solids and liquids are usually on a per-mass basis, and gases on a per-volume basis.
A specified amount of heating cannot be met with the amount of fuel calculated using only the HHV. There will be heat losses, the flue gas temperature will be greater than 60 ⁰F, and water in the flue gas will typically be vapor. (Seider 608)

==Hot Oil/Specialized Heat Transfer Fluids==

Specialized heat transfer fluids and hot oil circuits are used as heat sources when steam and fired heat is not appropriate. Specialized heat transfer fluids and hot oil circuits are extremely versatile in that they can be used in the temperature range of 323K to 673K. This range however is quite variable. For hot oils, the upper temperature limit is gauged based off of the thermal decomposition of the oil and coking/fouling of heat exchanger tubes.

Hot oil circuit systems are most commonly used when the plant has several small temperature heating requirements because it is more economically sound. Rather than having several fired heaters heat each small temperature requirement, it is much more economical to have one fired heater heat the hot oil and circulate that oil through each of the process to meet all of the heat needs. Hot oil systems are also generally favored over high pressure steam in processes that involve high pressure differentials between the process stream and high pressure steam. Hot oil systems are favored in this scenario because of safety concerns. If the steam were to leak, the pressure drop could cause serious safety issues.

Mineral oils are the most commonly used heat transfer fluids, and one prominent example is Dowtherm A. Dowtherm A is a combination of 26.5 wt% diphenyl in diphenyl oxide (Towler 2012) and is extremely thermally stable. These mineral oil systems generally require high flow rates.

=Process Cooling=

==Cooling Water==
Cooling water is used to cool and/or condense streams. Cooling water is usually circulated between process heat exchangers and a cooling tower. Water is cooled during downward motion by contact with air blown upwards, which can bring the water temperature to come within ~ 5 ⁰F of air’s wet-bulb temperature.Approximately 80% of the temperature reduction is due to evaporation of the cooling water and heat transfer to the surrounding air. Water can also be cooled in spray ponds and cooling ponds. Both work by providing high area for water to exchange heat with air.
Water in cooling towers is lost through drift and blowdown, and makeup is usually 1.5 to 3% of the circulating rate.
If a large natural body of water is nearby, it can be used as a source of cooling water and discharged downstream. This water is usually filtered to remove salts and impurities that may lead to fouling, but it is not treated.

==Refrigeration==
Cooling water can usually be used to cool a stream to ~ 100 ⁰F. Air can only cool to ~ 120 ⁰F. Air may be used in places where water is scarce or more costly to transport. To cool or condense streams below 100 ⁰F, a refrigerant is typically used. Chilled brine can also be used, but is less common.

Until 1995, CFC Freon R-12 (dichlorodifuloromethane) and HCFC Freon R-22 (chlorodifuloromethane) were commonly used refrigerants. However, the chlorine atom in the molecules caused the depletion of the ozone layer. The U.S. Clean Air Act Amendments of 1990 went into effect in 1995, and the production of these refrigerants has since stopped or been greatly reduced.

Cost estimates are based on ton-day of refrigeration, where a ton is the heat that needs to be removed to freeze 1 ton per day of water at 32 ⁰F. Substitutes have since been developed. R-134a is often used in place of R-12. According to the EPA, R-134a is not combustible at ambient conditions, and is essentially non-toxic under 400 ppm, and is not ozone-depleting. (Seider pg 607)

=Energy Efficiency=
One of the chief concerns in selecting and designing process utility systems for heating and cooling is how to achieve the most energy efficient design. There are countless means by which plants lose energy, two of the foremost being through the mixing of different temperature or pressure streams and through the disposal of warmed cooling water. (Seider, 2009) Proper utilities design can help mitigate each of these losses as well as many others. The energy efficiency of a plant will depend primarily on the heating and cooling methods that are being used and the overall system design itself. These two parameters are important in determining how well energy is transferred to the desired media as well as how well that energy is recovered and recycled.

==Hot Utility Efficiency==
As mentioned above, the most commonly used utilities for process heating in large scale processes are steam, fired heat, and hot oil heaters. Of these, steam is the most commonly used. Electricity, while efficient at creating power, is not a viable source of heat in large industrial processes. Common ranges of heating efficiency for these three methods are displayed in Table 1. (Towler and Sinnott, 2012; Broughton, 1994)

{| class="wikitable"
|+Table 1: Process Heating Efficiencies
|-
! Process Heating Method
! Typical Efficiency
|-
| Steam (via coal boiler)
| 72%
|-
| Steam (via gas boiler)
| 73%
|-
| Steam (via oil boiler)
| 75%
|-
| Fired Heat w/ Convective Section
| 85%
|-
| Fired Heat w/o Convective Section
| 60%
|-
| Hot Oil Heaters/Vaporizers
| 80-85%
|}

===Steam===
As steam is so popular for heating purposes, it is useful to analyze the numerous ways in which losses can occur in steam systems. There are five primary sources of inefficiency and heat loss in the generation and distribution of steam throughout a process plant:

*The heat content of boiler exhaust gas
*Incomplete combustion of boiler fuel
*Radiant losses from the boiler exterior
*Blowdown losses
*Distribution losses (pipe transport, steam traps, etc.)

The first four of these losses take place at the boiler and contribute to the heating efficiencies seen in Table 1 for steam created with coal, gas, and oil. Several methods can be used to minimize these losses, one of the foremost being the control of air-to-fuel ratio in the boiler. This ratio must be managed by weighing losses due to uncombusted fuel against losses due to the heat content of excess exhaust gas. (Broughton, 1994)

[[File:Air-fuel ratio.PNG|thumb|border|center|middle|upright=4|link=|atl=|Air-to-fuel ratio relationship with heat losses (Broughton Fig 2.3)]]

The air-to-fuel ratio can be optimized using a feedback process controller. The control system will analyze the oxygen content of exhaust air and adjust the incoming air flow rate to achieve a set percentage of excess air. While desired excess oxygen will vary depending on the type of fuel, it is consistently seen that in the minimum loss-range a 1% increase in excess air will result in a 1% decrease in efficiency. ("Oxygen Control")

In addition to air-to-fuel ratio management, steam heat losses at the boiler can be mitigated via energy recovery, which is discussed in further detail [[Utility_systems#Energy_Recovery|below]].

Outside the boiler, losses in distribution of steam throughout facilities can decrease energy efficiency by a significant amount. This can cause up to a 60% increase in losses, but typically results in an overall steam efficiency of 50-55% (down from approximately 75% at the boiler). (Broughton, 1994) There are several ways that this issue can be addressed. First and foremost are steam metering systems, which can be used to monitor heat supply to separate sections of the process facility, check efficiency of fuel use, and determine which processes in a given plant have the highest steam requirements. Another method of minimizing transport losses is to decentralize steam generation systems. It can be advantageous to have numerous smaller boilers rather than a single centralized boiler so that steam does not have to travel as long of distances to reach its destination.

==Cold Utility Efficiency==
Efficiency in cooling processes depends more on the method used, and by extension the amount of coolant needed. Water and air utility efficiencies depend primarily on the fluid flow required to maintain the system at a desired temperature, while powered refrigeration utilities (for colder processes) have efficiencies at approximately 60%--but ranging up to 90%--of Carnot cycle efficiency, a metric of ideal refrigeration efficiency. (Towler and Sinnott, 2012) Cooling systems represent by definition a loss of energy from the main process to the utility stream, and as such it is often useful to find other uses for the heated media before discharge.

==Energy Recovery==
Recovery and recycle of energy is perhaps the most important aspect of creating an energy efficient plant design, and it is important for process engineers to fully consider possibilities for heat recovery in order to aid in economic viability.

===Process Heat Exchange===
Heat exchanger networks are a very common energy recovery method in industrial processes. These networks most frequently allow energy from heated product streams to be transferred to feed streams that must be brought up to process temperature. (Biegler, 1997) More information on the function and design of heat exchanger networks can be found on the [[Heat_exchanger#Heat_Exchanger_Networks|heat exchanger wiki page]]. The following are several examples of energy recovery via heat exchange that are used in industrial processes.

In distillation columns the bottoms and distillate effluents have the potential for energy exchange. Though the condenser at the top of the column cannot supply its waste heat to the reboiler due to their respective temperatures, the effluent streams can supply heat to the feed via a feed-effluent exchanger. This reduces the utility requirements to raise the feed to column temperature. (Biegler, 1997)

[[File:Feed_sterilization.jpg|thumb|border|right|middle|upright=4|link=|atl=|Feed sterilization schematic. (Towler and Sinnott Fig. 3.30)]]

Feed sterilization, commonly used in the food industry, is a common application for heat recovery through process stream heat exchange. In this application, the feed must be heated for a certain amount of time to kill any biological contaminants, after which it can be used to heat the new raw feed for sterilization. This reduces energy demands on the steam heater and thus reduces cost. (Towler and Sinnott, 2012)

In multi-vessel batch processes it can be advantageous to exchange heat as the process fluid is being transferred between vessels. Like the previous examples, this reduces the utility needed to bring the colder feed up to process temperature, thus reducing costs. (Towler and Sinnott, 2012)

[[File:Batch_heat_exchange.jpg|thumb|border|center|middle|upright=4|link=|atl=|Multi-vessel batch heat exchange schematic (Towler and Sinnott Fig. 3.31)]]

===Utility Regeneration===
[[File:Waste_heat_boiler.jpg|thumb|border|right|middle|200px|upright=4|link=|atl=|Industrial modular HRSG]]
When recovery of waste heat via transfer to other process streams is inconvenient or impossible, energy efficiency can still be improved through the regeneration of utilities. This is commonly done through the regeneration of steam by removing heat from exiting streams or from highly exothermic reactions. Waste heat in exiting streams can be removed via heat recovery steam generators (HRSGs), and is most often used on exiting gas streams. Heat recovery from reactions is a viable option when the reactor temperature will be at 150 C or above, as this will create steam at high enough pressure to be used in other processes. (Towler and Sinnott, 2012)

In the case of steam, waste heat and water treatment losses can be recovered from the utility generation process itself. One of the most common ways to do this is with an economizer. As seen in the schematic below, economizers have heat exchangers fit to the exhaust gas flow in order to transfer waste heat from these gases to the incoming boiler feed water. This can result in fuel energy savings of approximately 15% for typical excess air percentages in steam boilers. (Broughton, 1994; "Economizers")

[[File:Economizer.png|thumb|border|center|middle|upright=4|link=|atl=|Economizer schematic ("Economizers") and economizer energy recovery correlations (Broughton Fig 2.2)]]

[[File:Condensate recycle.PNG|thumb|border|right|middle|200px|upright=4|link=|atl=|Condensate return fuel savings (Broughton Fig 2.4)]]

Additionally, there are still more opportunities for recovery in steam generation processes. One typical case is the use of waste heat to preheat air entering the boiler. Optimistically, this can result in a 5% improvement in heat recovery for the system. Another method is the recycle of steam condensate back to the boiler. This has two main benefits: it retains the heat still present in the condensed steam and it retains the pretreatment chemicals added to the steam. (Broughton, 1994)

Furthermore, there is opportunity for energy recovery in the expansion of compressed gas through a turbine to create electricity, a process that can be economically viable given sufficiently high flows or pressure. One of the primary application of this type of energy recovery is in the creation of medium pressure and low pressure steam. In most processes, all steam is generated as high pressure steam and can be expanded through a turbine to decrease its pressure. Such technology has also been used in processes to synthesize ammonia, perform air separations, and synthesize nitric acid. (Towler and Sinnott, 2012) Recently, however, there has been a particularly strong interest for energy recovery in the natural gas industry, when gas is decompressed from major pipelines to residential low-pressure pipelines. A 2001 study estimated that there is the potential to recover 21 TWh, representing 11% of natural gas transport energy, via gas expansion. (Lehman)

=Process water and boiler-feed water=
Process water is water that will be directly used in the process. Boiler-feed water (BFW) is used to produce steam. Both may need to be purified to prevent impurities from contaminating a process or from foul equipment. It can be used as a cooling stream when the temperature of the stream to be cooled is greater than ~300 ⁰F. Cost of BFW can be partially offset by the steam credit.

Process water that undergoes moderate pretreatment can cost ~ $0.75/1,000 gal.

Extensive treatment ~ $6.00/1,000 gal.

Sterilized for pharmaceutical processes ~ $550/1,000 gal. (Seider pg 608)

=Demineralized Water=
In demineralized water, minerals have been removed by ion exchange. In boiler feed water, this is to prevent salt deposition, corrosion, formation of foam, and sluicing. In process water, the ions may contaminate the process.

=Waste Treatment=
Most chemical processes will produce some sort of waste. Disposal occurs to the atmosphere (in the case of some gases), sewers, body of water, or a landfill. Waste may require some treatment before disposal to meet regulations. Depending on process economics, byproducts may be recovered and processed. (Seider 2009 pg 609)

==Wastewater Treatment==
The United States EPA regulates industrial wastewater disposal through the Clean Water Act, introduced in 1948 as the Federal Water Pollution Control Act and amended to its current form in 1972. The sweeping 1972 amendments allowed the EPA to prevent industries and persons from discharging contaminated water into fresh water sources and set water quality standards. (Summary of the Clean Water Act) In accordance with this law, process plants in the United States treat wastewater at on-site or near-site treatment centers before releasing it into the surrounding environment.

Wastewater effluent streams, along with water runoff from around the plant, are treated to control for pH, toxicity, suspended solids, and biological oxygen demand (for aquatic life protection) prior to discharge. Each of these controls is typically addressed with a separate method. Acidity and basicity is balanced through the addition of an acid or alkaline solution. Toxic wastewater may be treated with chemical processes or simply diluted to safe concentrations. Suspended solids can be removed via filtration and/or with clarifiers. Oxygen demand of wastewater can be mitigated using activated sludge treatment processes. Once the water quality complies with the EPA, and state-mandated, regulations, it can be safely released. More information on the large number of industry-specific guidelines for waste effluent can be found on the EPA website (http://www.epa.gov/eg/industrial-effluent-guidelines).

=Air-Pollution Management=
===Introduction===

In the United States air pollution is regulated in the Clean Air Act, and almost all pollutant emitting plants are regulated under this law. The types of plants that can release significant emissions include petroleum refineries, sulfur recovery plants, carbon-black plants, fuel conversion plants, chemical process plants, fossil fuel plants, and petroleum storage and transfer facilities. To receive permission to construct a plant must undergo a review to show that it will not cause a violation of the Ambient Air Quality Standards(Peters, 2003).

==Methods==

There are two major types of pollutants that are released into the air, particulates and and gaseous pollutants. Particulates can be removed with mechanical forces while gaseous pollutants typically need to removed by chemical or physical means (Peters, 2003).

Particulates and volatile pollutants that need to be removed before disposal may be present. Particle removal equipment includes: cyclones, wet scrubbers, electrostatic precipitators, and fabric-filter systems (Seider 2009) such as bag filters (. The two charts below are from ''Plant Design and Economics for Chemical Engineers'' and show the types of equipment, separation methods, and particle sizes in different pollutant separation technologies (Peters, 1991).

[[File:AirPolChart.png|thumb|border|center|middle|upright=4|link=|atl=|Different Methods of Particle Separation, the Particle Sizes they Can Remove, and the Technologies Used]]

[[File:AirPolTable.png|thumb|border|center|middle|upright=4|link=|atl=| This Chart Shows the Specifications and Limitations of Different Separations Technologies Including Particle Size, Efficiencies, and Temperatures(Peters, 1991)]]

Methods for removing inorganic and organic gaseous pollutants include: absorption, adsorption, condensation, and combustion (Seider 2009 pg 609). A list of typical of gases pollutants and their sources from ''Plant Design and Economics for Chemical Engineers'' is shown below (Peters, 1991).

[[File:GaseousPollutants.png|thumb|border|center|middle|upright=4|link=|atl=| Common Gaseous Pollutants and their Sources]]

Typically Gas-liquid absorption processes are done completed in a vertical, countercurrent, flow through packed, plate, or spray towers. These systems require good liquid-gas contact and proper equipment. These systems also often have significant energy consumption because of large pressure drops (Peters, 2003). For high volume systems absorption by scrubbing with water or another solvent is the most widely used method (Towler, 2012). Dry adsorbents can be used to remove the last races of gaseous pollutants. Adsorption typically requires blowers, condensers, separators, and controls. You also typically need two packed beds so that one can be used while the other is regenerated. Examples of adsorbents are molecular sieves and activated carbon. Incineration is typically used when there are gas streams that have no recovery value. This can be done with direct flame or catalytic oxidation. Catalytic oxidation usually has higher capital costs, but lower operating costs because it does not require fuel.
===Wet Scrubbers===
Wet scrubber use lime or limestone and water to remove SO2 and acid gases. The mixture can be injected into a scrubber or the gases can be bubbled through this mixture. This results in removal of 90-98% of SO2 and and acid gases (Clean Coal Technologies).

===Dry Scrubbers===
Dry scrubber blow powdered adsorbents into a vessel with gases and then after it has captured the SO2 and acid gases it is separated from gas using a fabric filter. These systems remove 90-93% of the contaminants (Clean Coal Technologies).

===Low-NOx Burners===
The purpose of low NOx burners is to decrease the amount of NOx created when the coal is burned. This is done by injecting coal and air in boilers. This can result in 40-50% NOx reduction. If air is injected into the area above the burner this can actually cause almost 70% NOx reduction (Clean Coal Technologies).

===Selective Non-Catalytic and Catalytic Reductions===
These systems inject ammonia into gases to remove NOx. The Catalytic reductions add a catalyst to the ammonia being injected to aid in the removal. Con-catalytic reductions result in about 35% removal, but adding a catalyst can increase that amount to about 90%. The catalytic-reduction can remove up to 80% of mercury as well (Clean Coal Technologies).

===Fabric Filters===
Also known as baghouses,these [[Separation processes#Filtration| filters]] remove particulates by passing air through filters. These can separate as much as 99.9% of particulate matter (Clean Coal Technologies).

===Electrostatic Precipitators===
Electrostatic Precicpitators remove particulate matter as gas passes through a device that has charged metal plates. The particles are then removed because of static electricity. These systems can remove between 99-99.9% of particulate matter (Clean Coal Technologies).

===Super-critical Boilers===
Super-Critical Boilers and Ultra-Supercritical Boilers operate at temperature and pressure higher than regular boilers. By operating at higher temperature these systems become more efficient. Super Critical Boilers typically have 10%-20% CO2 emissions than other similar sub-critical technologies. Ultra-Supercritical boilers can be as much as 30% more efficent than sub-critical technologies (Clean Coal Technologies).

===More Information===

For more information regarding some of these separations equipment see [[Solids-involved equipment]]

For more information Regarding Cyclones see [[Separation processes#Cyclones| Cyclones]], and for modeling cyclones see their[[Solids-involved equipment#HYSYS Simulation|HYSYS Simulation]]

You can find more information on [[Separation processes#Absorption| Absorption]] and [[Separation processes#Adsorption| Adsorption]] in [[Separation processes]]

===Outcomes===

====United States Common Pollutant Emissions====

The United States implemented the Clean Air Act in 1970 and since then emissions in the U.S. have been drastically reduced. As a matter of fact despite increases in the population in the last 35 years the amount of pollutants emitted have decreased by almost 70% (EPA). More importantly this demonstrates that reasonable efforts can be put towards environmental protection without causing too much harm to industry.

[[File:EPA.png|thumb|border|center|middle|upright=4|link=|atl=| GDP and Other Growth Factors Vs Common Pollutants Released]]

===China Clean Coal===

====Success Through 2005====

Coal is a very inexpensive and abundant source of energy and is abundant in China (Xu 2010). In China Coal the cause of 90% of SO2 emissions, 70% of dust emissions, and 67% of NOx emissions, and 70% of CO2 emissions. These numbers are staggering considering that Coal has never been more than 50 percent of China's Energy Supply (Xu 2010).

[[File:Coal.jpg|thumb|border|center|middle|upright=4|link=|atl=| Coal Consumption for Power and the Percentage of Coal Used (Xu 2010)]]

Despite the increasing coal consumption, high efficiency electric dust removal systems with efficiencies that are as high as 99.6% have greatly decreased soot emissions to 32% below 1980's levels as of 2005.

[[File:Soot.jpg|thumb|border|center|middle|upright=4|link=|atl=| SO2 and Soot Emissions from 1981 to 2005 (Xu 2010)]]

====Technologies and Life Cycle Comparison====
As of 2010 China was consuming 48.2% of coal globally. Four potential technologies integrated gasification combined cycle (IGCC), sub-critical coal power generation (Sub-C); super-critical coal power generation (Super-C) ultrasuper-critical coal power generation (USC). These technologies are compared mostly on the basis of net generating efficiency and efficiency. Net generating efficiency is the output of the plant divided by the total available energy in the fuel used (Liang, 2013).

where efficiency is define as: η=E/Etotal

and

Etotal=Emining+Etransportation+Egeneration

[[File:technologies.PNG|thumb|border|center|middle|upright=4|link=|atl=|A comparison of many of the parameters important to lfe-cycle analysis (Liang, 2013)]]

This demonstrates that USC has the best net generation efficiency, while also having the largest capacity for a single system.

[[File:Lifecycle2.PNG|thumb|border|center|middle|upright=4|link=|atl=| (Liang, 2013)]]

This breaks down the energy efficiency of each system and demonstrates that USC is the most efficient

[[File:Lifecycle.PNG|thumb|border|center|middle|upright=4|link=|atl=| (Liang, 2013)]]

IGCC and USC are lowest on different element, but since these all have different global warming potential it is difficult to tell which is the most efficient.

[[File:GWP.PNG|thumb|border|center|middle|upright=4|link=|atl=| Environemntal measures and their explanations (Liang, 2013)]]

This table breaks down the different measures of which systems are better

[[File:Measures.PNG|thumb|border|center|middle|upright=4|link=|atl=| Environemntal measures and their explanations (Liang, 2013)]]

And the table above defines all the relevant terms.

[[File:Price.PNG|thumb|border|center|middle|upright=4|link=|atl=| Capital cost in $/kW of clean coal power-generation technologies (Liang, 2013)]]

This demonstrates that the capital cost per unit of energy production is very high for IGCC, but shows that Super-C and USC are very competitive.

And the table below defines all the relevant terms.

[[File:Measures.PNG|thumb|border|center|middle|upright=4|link=|atl=| Environemntal measures and their explanations (Liang, 2013)]]

All of the data suggests that USC is the highest in both energy efficiency and net generating efficiency while competitive in price and the lowest on emissions among the different boilers.

=Solid Waste=
U.S. federal regulations require that solid waste be classified as hazardous or nonhazardous. Conditions for a classification of hazardous include: ignitability, corrosivity, reactivity, toxicity, or posing a substantial threat to the surrounding environment and its inhabitants. Hazardous waste must be treated on- or near-site before being removed in containers. Non-hazardous waste may be landfilled or incinerated in some cases. A typical estimate of costs for waste disposal is $0.03/lb for nonhazardous solids and $0.10/lb for hazardous solids. (Seider 2009 pg 609)

=References=

*Biegler LT, Grossmann IE, Westerberg AW. ''Systematic Methods of Chemical Process Design'' Prentice-Hall: Upper Saddle River, 1997.

*Broughton, J. ''Process Utility Systems: Introduction to Design, Operation and Maintenance'' Institution of Chemical Engineers: Rugby, Warwickshire, UK, 1994.

*Clean Coal Technologies. America's Coalition for Clean Coal Electricity. http://www.americaspower.org/clean-coal-technologies-1663/

*Duke Energy Company (2013). How Do Coal Fired Plants Work? Charlotte: Duke Energy.

*"Economizers." Energy Solutions Center Inc.: Boiler Consortium. http://www.cleanboiler.org/Eff_Improve/Efficiency/Economizers.asp. Accessed February 21, 2016.

*Lehman B, Worrell E. ''Electricity Production from Natural Gas Pressure Recovery Using Expansion Turbines.'' Lawrence Berkeley National Laboratory; 2001.

*"Overview of the Clean Air Act and Air Pollution." Environmental Protection Agency. November 17, 2015. http://www.epa.gov/clean-air-act-overview. Accessed February 5, 2016.

*"Oxygen Control." Energy Solutions Center Inc.: Boiler Consortium. http://www.cleanboiler.org/Eff_Improve/Efficiency/Oxygen_Control.asp. Accessed February 21, 2016.

*Peters, Max S.; Timmerhaus, Klaus D.; West, Ronald E. (2003). " Plant Design and Economics for Chemical Engineers." McGraw Hill Higher Education.

*Ruina Xu Energy (2010). "Clean coal technology development in China." Elsevier.

*Seider, Seader, Lewin, Widagdo. (2009). ''Plant Design and Economics for Chemical Engineers, 5th Edition.'' Hoboken: Wiley.

*Seider, Seader, Lewin. (2008). ''Product and Process Design Principles, 2nd Edition.'' Hoboken: Wiley.

*Summary of the Clean Water Act. United States EPA website. http://www.epa.gov/laws-regulations/summary-clean-water-act

*Towler, G.P. and Sinnot, R. (2012). ''Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design.'' Elsevier.

*Turton R.; Bailie, R.C.; Whiting, W.B.; Shaeiwitz J.A.; Bhattacharyya D. (2012). ''Analysis, Synthesis, and Design of Chemical Processes.'' Upper Saddle River: Prentice Hall.

*G.D. Ulrich, A Guide to Chemical Engineering Process Design and Economics, Wiley: New York, 1984.

*Xiaoye Liang, Zhihua Wang, Zhijun Zhou, , Zhenyu Huang, Junhu Zhou, and Kefa Cen. (2013). "Up-to-date life cycle assessment and comparison study of clean coal power generation technologies in China" Elsevier.

Utility systems

2016-02-22T05:08:27Z

Bts511: /* Technologies and Life Cycle Comparison */

Authors: David Chen, [2014] Joshua Lee, [2015] Brett Sleyster, [2016] and Tom Aunins [2016] 

Stewards: David Chen, Jian Gong, and Fengqi You

Date Presented: January 13, 2014 /Date Revised: January 14, 2014

 


__TOC__

=Introduction=
Many chemical processes do not take place at ambient temperature or pressures. In order to reach these non-ambient conditions, utilities will have to be used to raise or lower temperatures and compress gases. (Towler, Towler/UOP) Utilities often contribute 5 to 10% of the price of a product, and may come from public or private utility companies or on-site plants. For purchased utilities, costs depend on consumption, while for company-owned utilities, there will be both capital and operating costs. They include things such as steam for heating, electricity, cooling water, refrigeration, fuels such as natural gas, wastewater treatment, waste disposal, and landfill. Steam is often the largest utility cost. Cogeneration unit can supply electricity accompanied with different steam pressures. (Seider 2009)

=Electricity=

Electricity is used to power many different kinds of equipment. It has many advantages: it is efficient (> 90%), reliable, available in a wide range of power, shaft speeds, designs, lifetimes, convenience, costs, and maintenance. It is generally used up to 200 hp, and sometimes over 10,000 Hp. In chemical process plants, the electricity demand is generally determined by the work or energy required for compression, pumping, air cooling, lights, and many other items. This electricity often times is purchased from local electricity providers, but many plants generate their own electricity via sophisticated processes.

Electricity is rarely used as a primary heat utility in large-scale chemical plants for a variety of reasons. The main disadvantages of using electricity as a heat utility are as follows (Towler 2012):

:*Heat from electricity is two to three times more expensive than heat from fuels. This is attributed to the lack of efficiency when creating heat from electricity.

:*Electrical heating units are expensive, require high maintenance, and must comply with strict safety regulations.

:*Electrical heating units are unsafe compared to steam heating units. In steam systems, the physically steam controls the temperature, whereas in electrical heating units temperature is controlled by temperature controllers, which can fail or burn out.

The use of electricity carries with it some hazards depending on the environment. Extra care must be taken when using electrically-powered equipment in areas which may have combustible fluids, vapors, or dust, and where liquids may be present. (Seider 2009 pg 606)

==Conventional Power Station==

In general, most electricity is generated from a conventional coal-fired process, whether it be on-site or purchased from a provider. Coal-fired processes have been used to create electricity throughout history, and technological advances have increased its efficiency and use worldwide. In a coal-fired steam station—much like a nuclear station—water is turned into steam, which in turn drives turbine generators to produce electricity. There are several variations on how to create energy from coal, but here are the basics of how a coal-fired process works:

[[File:coalfired.png|thumb|border|center|middle|upright=4|link=|atl=|General Coal-Fired Process Diagram (Duke Energy Company)]]

*Heat is created:
::Before the coal is burned, it is pulverized to the fineness of talcum powder. It is then mixed with hot air and blown into the firebox of the boiler. Burning in suspension, the coal/air mixture provides the most complete combustion and maximum heat possible.

*Water turns to steam:
::Highly purified water, pumped through pipes inside the boiler, is turned into steam by the heat. The steam reaches temperatures of up to 1,000 degrees Fahrenheit and pressures up to 3,500 pounds per square inch, and is piped to the turbine.

*Steam turns the turbine:
::The enormous pressure of the steam pushing against a series of giant turbine blades turns the turbine shaft. The turbine shaft is connected to the shaft of the generator, where magnets spin within wire coils to produce electricity.

*Steam is converted back to water:
::After doing its work in the turbine, the steam is drawn into a condenser, a large chamber in the basement of the power plant. In this important step, millions of gallons of cool water from a nearby source (such as a river or lake) are pumped through a network of tubes running through the condenser. The cool water in the tubes converts the steam back into water that can be used over and over again in the plant.

*Repeat:
::The cooling water is returned to its source without any contamination, and the steam water is returned to the boiler to repeat the cycle.

===Advantages & Disadvantages of Coal-Fired Energy Production===

[[File:coalfiredtable.png|thumb|border|center|middle|upright=4|link=|atl=|Advantages and Disadvantages of Coal-Fired Processes (Seider 2008)]]

==Gas-Turbine Cogeneration Process==

When generating energy on-site, many plants use a gas-turbine cogeneration process. The thermal efficiency of a gas-turbine process is in the range of 70-80% while conventional power stations, such as coal-fired processes, have a 30-40% efficiency. The lower efficiency in more conventional power stations is attributed to wasted heat in the exhaust steam in the condenser. One example of a gas-turbine process is outlined in the following figure. Figure 3.1 is a gas-turbine cogeneration process with a heat recovery steam generator (waste-heat) boiler.

[[File:cogeneration.png|thumb|border|center|middle|upright=4|link=|atl=|Gas-Turbine Cogeneration with a heat recovery steam generator boiler (Towler 2012 Fig 3.1)]]

Overall, the process illustrated is not much different from a coal-fired process. The main differences are that the cogeneration process creates both electricity and a heat utility, and cogeneration processes use natural gas instead of coal. Many of the advantages and disadvantages are similar to those of the coal-fired process, but the cogeneration has a much higher efficiency, creates heat to be used in another process, and uses a more volatile and expensive fuel. The main advantage of cogeneration over coal-fired energy production is that heat is not wasted. In coal-fired processes, heat is released and wasted during electricity generation. Cogeneration captures some, if not all of the byproduct for heat, which is an extremely useful utility that will be discussed in subsequent sections. In summary, the cogeneration plant is superior to the coal-fired process because of its higher efficiency and ability to create a useful heat utility.

Obviously any engineer would design the cogeneration plant to meet at least the energy requirement necessary for plant operation, but cogeneration plants often times are designed to exceed the plant electricity requirement to drive another source of capital. Many describe this scenario as a "make or buy" scenario (Towler 2012). This scenario provides chemical producers leverage when negotiating contracts with outsourced electricity providers and this allows plants to purchase electricity at a wholesale price. This is a huge advantage of considering on-site electricity production because electricity is needed in relatively high quantities for all chemical plants. Being able to minimize electricity costs, or even profit off of electricity production is a huge economical consideration that all plants employ.

=Process Heating=
The key objective of this section is to discuss how processes are heated. Heating utilities are necessary for proper usage of distillers, reactors, condensers, and several other integral types of equipment. More specifically, steam, fired heat, and hot oil/specialized heat transfer fluids will be discussed in the following subsections.

==Steam==
Steam is the most commonly used heat utility used in chemical plants, and as a result understanding how it is used is essential in the study of Utility systems. Steam is used both as a process fluid (feedstock, diluent to absorb heat of reaction, heating agent, and stripping agent in absorbers and adsorbers ) and utility. It can be used to drive pumps and compressors, ejectors (for producing a vacuum), and heat exchangers. As one can clearly see, steam is a versatile, and useful utility.

Here are a few advantages of using steam as opposed to other methods of process heating (Towler 2012):

:*By controlling the pressure of the steam, one can control the temperature at which the heat is released. Having a strong control over the temperature is essential in several processes. =
:*Steam is an efficient heat source because the heat of condensation of steam is very high. Meaning that there is is high output per mass of utility at a constant temperature.
:*Heat exchangers that use steam are relatively cheap because condensing steam has a high heat transfer coefficient.
:*Steam is nonflammable, nontoxic, and inert to several process fluids.

Chemical plants generally have a network of pipelines exclusively for providing steam. These networks generally have steam at a low pressure, a medium pressure, and a high pressure. The image below illustrates a basic steam system.

[[File:Steam.png|thumb|border|center|middle|upright=4|link=|atl=|(Towler 2012 Fig 3.2)]]

In the diagram above, boiler feed water at a high pressure is preheated and fed to other boilers. These other boilers superheat the steam to create a high pressure and high temperature steam stream. The steam is superheated past the dew point to account for heat loss in the pipelines. A portion of the high pressure steam is used for process heating in areas of the plant that require high temperatures. The rest of the high pressure steam is turned into medium pressure steam by valves and steam turbines. The medium pressure steam is then used to heat medium temperature processes and to form low pressure steam. The low pressure steam can be used to heat low pressure processes and it can be expanded in condensing turbines to create shaft work and energy. In summary, steam can be used for an innumerable amount of action items in a plant. High pressure, medium pressure, and low pressure steam can all be used as a heat source. Low pressure steam has utilities in creating electricity and it also has several other uses.

==Fired Heat==

In many cases, processes in a plant require a heat source stronger than high pressure and temperature steam. That is when fired heat is used, which is generally at temperatures above 523K. Streams can be heated directly in the furnace tubes or via a hot oil circuit or heat transfer fluid, which will be discussed in detail in the next section. Most fired heaters use natural gas as fuel because it burns cleaner than fuel oil. A cleaner burning fuel is always advantageous due to environmental and safety concerns. Furthermore, natural gases usually result in less wear and tear in burners and fuel lines.

Depending on the application of the fired-heater, different design specifications can be implemented to make the fired-heater as efficient as possible. The basic construction of a fired heater starts with a cylindrical chamber that is lined with with refractory bricks. The stream that is to be heated flows through tubes inside of the furnace. These tubes can be arranged in several different arrangements such as, around the walls of the furnace, or in horizontal or vertical banks (Towler 2012). The figure below illustrates the basic construction of the fired-heater and varying tube arrangements.

[[File:firedheater.png|thumb|border|center|middle|upright=4|link=|atl=|Basic Fired-Heater Diagrams (Towler 2012 Figure 19.66)]]

Fuel is burned to heat the entire furnace, and the heat transfer occurs from the combustion gases inside of the furnace across the tubes that are filled with our desired stream. The heat transfer between the tubes and the furnace is accomplished primarily via radiation. Modern designs take advantage of convective heat transfer by adding a smaller section on top of the fired-heater where the combustion gases flow over banks of tubes as seen in (c) in the figure above. Heat transfer can be further improved via convection by adding fins or pins in the combustion section.

The cost of fired heat can be calculated by the cost of fuel fired. Natural gas and heating oil are traded as commodities and prices can be found at many online trading sites or business news sites (i.e., www.cnn.money.com). Past historic prices for forecasting can be found in the Oil and Gas Journal or from the U.S. Energy Information Adminstration (www.eia.gov).

===Fuels===
Fuel is burned in utility facilities such as boilers, electricity generation, and cogeneration, and can be in solid, liquid, or gas form. It can also be burned to provide heating for a process or stream or to drive pumps and compressors. The fuel is usually burned with excess air to ensure complete combustion.

A way of quantifying the amount of heat generated is by using the heating values. Higher heating value (HHV) and the lower heating value (LHV) are used. The heating is the total heat evolved by complete combustion of a fuel with dry air with both at 60 ⁰F and the flue gas after combustion brought back down to ⁰F. If the water vapor in the flue gas is not condensed, we obtain the LHV. If the water vapor is condensed, the value of heat evolved is a bit higher, and this is the HHV.
Heating values for solids and liquids are usually on a per-mass basis, and gases on a per-volume basis.
A specified amount of heating cannot be met with the amount of fuel calculated using only the HHV. There will be heat losses, the flue gas temperature will be greater than 60 ⁰F, and water in the flue gas will typically be vapor. (Seider 608)

==Hot Oil/Specialized Heat Transfer Fluids==

Specialized heat transfer fluids and hot oil circuits are used as heat sources when steam and fired heat is not appropriate. Specialized heat transfer fluids and hot oil circuits are extremely versatile in that they can be used in the temperature range of 323K to 673K. This range however is quite variable. For hot oils, the upper temperature limit is gauged based off of the thermal decomposition of the oil and coking/fouling of heat exchanger tubes.

Hot oil circuit systems are most commonly used when the plant has several small temperature heating requirements because it is more economically sound. Rather than having several fired heaters heat each small temperature requirement, it is much more economical to have one fired heater heat the hot oil and circulate that oil through each of the process to meet all of the heat needs. Hot oil systems are also generally favored over high pressure steam in processes that involve high pressure differentials between the process stream and high pressure steam. Hot oil systems are favored in this scenario because of safety concerns. If the steam were to leak, the pressure drop could cause serious safety issues.

Mineral oils are the most commonly used heat transfer fluids, and one prominent example is Dowtherm A. Dowtherm A is a combination of 26.5 wt% diphenyl in diphenyl oxide (Towler 2012) and is extremely thermally stable. These mineral oil systems generally require high flow rates.

=Process Cooling=

==Cooling Water==
Cooling water is used to cool and/or condense streams. Cooling water is usually circulated between process heat exchangers and a cooling tower. Water is cooled during downward motion by contact with air blown upwards, which can bring the water temperature to come within ~ 5 ⁰F of air’s wet-bulb temperature.Approximately 80% of the temperature reduction is due to evaporation of the cooling water and heat transfer to the surrounding air. Water can also be cooled in spray ponds and cooling ponds. Both work by providing high area for water to exchange heat with air.
Water in cooling towers is lost through drift and blowdown, and makeup is usually 1.5 to 3% of the circulating rate.
If a large natural body of water is nearby, it can be used as a source of cooling water and discharged downstream. This water is usually filtered to remove salts and impurities that may lead to fouling, but it is not treated.

==Refrigeration==
Cooling water can usually be used to cool a stream to ~ 100 ⁰F. Air can only cool to ~ 120 ⁰F. Air may be used in places where water is scarce or more costly to transport. To cool or condense streams below 100 ⁰F, a refrigerant is typically used. Chilled brine can also be used, but is less common.

Until 1995, CFC Freon R-12 (dichlorodifuloromethane) and HCFC Freon R-22 (chlorodifuloromethane) were commonly used refrigerants. However, the chlorine atom in the molecules caused the depletion of the ozone layer. The U.S. Clean Air Act Amendments of 1990 went into effect in 1995, and the production of these refrigerants has since stopped or been greatly reduced.

Cost estimates are based on ton-day of refrigeration, where a ton is the heat that needs to be removed to freeze 1 ton per day of water at 32 ⁰F. Substitutes have since been developed. R-134a is often used in place of R-12. According to the EPA, R-134a is not combustible at ambient conditions, and is essentially non-toxic under 400 ppm, and is not ozone-depleting. (Seider pg 607)

=Energy Efficiency=
One of the chief concerns in selecting and designing process utility systems for heating and cooling is how to achieve the most energy efficient design. There are countless means by which plants lose energy, two of the foremost being through the mixing of different temperature or pressure streams and through the disposal of warmed cooling water. (Seider, 2009) Proper utilities design can help mitigate each of these losses as well as many others. The energy efficiency of a plant will depend primarily on the heating and cooling methods that are being used and the overall system design itself. These two parameters are important in determining how well energy is transferred to the desired media as well as how well that energy is recovered and recycled.

==Hot Utility Efficiency==
As mentioned above, the most commonly used utilities for process heating in large scale processes are steam, fired heat, and hot oil heaters. Of these, steam is the most commonly used. Electricity, while efficient at creating power, is not a viable source of heat in large industrial processes. Common ranges of heating efficiency for these three methods are displayed in Table 1. (Towler and Sinnott, 2012; Broughton, 1994)

{| class="wikitable"
|+Table 1: Process Heating Efficiencies
|-
! Process Heating Method
! Typical Efficiency
|-
| Steam (via coal boiler)
| 72%
|-
| Steam (via gas boiler)
| 73%
|-
| Steam (via oil boiler)
| 75%
|-
| Fired Heat w/ Convective Section
| 85%
|-
| Fired Heat w/o Convective Section
| 60%
|-
| Hot Oil Heaters/Vaporizers
| 80-85%
|}

===Steam===
As steam is so popular for heating purposes, it is useful to analyze the numerous ways in which losses can occur in steam systems. There are five primary sources of inefficiency and heat loss in the generation and distribution of steam throughout a process plant:

*The heat content of boiler exhaust gas
*Incomplete combustion of boiler fuel
*Radiant losses from the boiler exterior
*Blowdown losses
*Distribution losses (pipe transport, steam traps, etc.)

The first four of these losses take place at the boiler and contribute to the heating efficiencies seen in Table 1 for steam created with coal, gas, and oil. Several methods can be used to minimize these losses, one of the foremost being the control of air-to-fuel ratio in the boiler. This ratio must be managed by weighing losses due to uncombusted fuel against losses due to the heat content of excess exhaust gas. (Broughton, 1994)

[[File:Air-fuel ratio.PNG|thumb|border|center|middle|upright=4|link=|atl=|Air-to-fuel ratio relationship with heat losses (Broughton Fig 2.3)]]

The air-to-fuel ratio can be optimized using a feedback process controller. The control system will analyze the oxygen content of exhaust air and adjust the incoming air flow rate to achieve a set percentage of excess air. While desired excess oxygen will vary depending on the type of fuel, it is consistently seen that in the minimum loss-range a 1% increase in excess air will result in a 1% decrease in efficiency. ("Oxygen Control")

In addition to air-to-fuel ratio management, steam heat losses at the boiler can be mitigated via energy recovery, which is discussed in further detail [[Utility_systems#Energy_Recovery|below]].

Outside the boiler, losses in distribution of steam throughout facilities can decrease energy efficiency by a significant amount. This can cause up to a 60% increase in losses, but typically results in an overall steam efficiency of 50-55% (down from approximately 75% at the boiler). (Broughton, 1994) There are several ways that this issue can be addressed. First and foremost are steam metering systems, which can be used to monitor heat supply to separate sections of the process facility, check efficiency of fuel use, and determine which processes in a given plant have the highest steam requirements. Another method of minimizing transport losses is to decentralize steam generation systems. It can be advantageous to have numerous smaller boilers rather than a single centralized boiler so that steam does not have to travel as long of distances to reach its destination.

==Cold Utility Efficiency==
Efficiency in cooling processes depends more on the method used, and by extension the amount of coolant needed. Water and air utility efficiencies depend primarily on the fluid flow required to maintain the system at a desired temperature, while powered refrigeration utilities (for colder processes) have efficiencies at approximately 60%--but ranging up to 90%--of Carnot cycle efficiency, a metric of ideal refrigeration efficiency. (Towler and Sinnott, 2012) Cooling systems represent by definition a loss of energy from the main process to the utility stream, and as such it is often useful to find other uses for the heated media before discharge.

==Energy Recovery==
Recovery and recycle of energy is perhaps the most important aspect of creating an energy efficient plant design, and it is important for process engineers to fully consider possibilities for heat recovery in order to aid in economic viability.

===Process Heat Exchange===
Heat exchanger networks are a very common energy recovery method in industrial processes. These networks most frequently allow energy from heated product streams to be transferred to feed streams that must be brought up to process temperature. (Biegler, 1997) More information on the function and design of heat exchanger networks can be found on the [[Heat_exchanger#Heat_Exchanger_Networks|heat exchanger wiki page]]. The following are several examples of energy recovery via heat exchange that are used in industrial processes.

In distillation columns the bottoms and distillate effluents have the potential for energy exchange. Though the condenser at the top of the column cannot supply its waste heat to the reboiler due to their respective temperatures, the effluent streams can supply heat to the feed via a feed-effluent exchanger. This reduces the utility requirements to raise the feed to column temperature. (Biegler, 1997)

[[File:Feed_sterilization.jpg|thumb|border|right|middle|upright=4|link=|atl=|Feed sterilization schematic. (Towler and Sinnott Fig. 3.30)]]

Feed sterilization, commonly used in the food industry, is a common application for heat recovery through process stream heat exchange. In this application, the feed must be heated for a certain amount of time to kill any biological contaminants, after which it can be used to heat the new raw feed for sterilization. This reduces energy demands on the steam heater and thus reduces cost. (Towler and Sinnott, 2012)

In multi-vessel batch processes it can be advantageous to exchange heat as the process fluid is being transferred between vessels. Like the previous examples, this reduces the utility needed to bring the colder feed up to process temperature, thus reducing costs. (Towler and Sinnott, 2012)

[[File:Batch_heat_exchange.jpg|thumb|border|center|middle|upright=4|link=|atl=|Multi-vessel batch heat exchange schematic (Towler and Sinnott Fig. 3.31)]]

===Utility Regeneration===
[[File:Waste_heat_boiler.jpg|thumb|border|right|middle|200px|upright=4|link=|atl=|Industrial modular HRSG]]
When recovery of waste heat via transfer to other process streams is inconvenient or impossible, energy efficiency can still be improved through the regeneration of utilities. This is commonly done through the regeneration of steam by removing heat from exiting streams or from highly exothermic reactions. Waste heat in exiting streams can be removed via heat recovery steam generators (HRSGs), and is most often used on exiting gas streams. Heat recovery from reactions is a viable option when the reactor temperature will be at 150 C or above, as this will create steam at high enough pressure to be used in other processes. (Towler and Sinnott, 2012)

In the case of steam, waste heat and water treatment losses can be recovered from the utility generation process itself. One of the most common ways to do this is with an economizer. As seen in the schematic below, economizers have heat exchangers fit to the exhaust gas flow in order to transfer waste heat from these gases to the incoming boiler feed water. This can result in fuel energy savings of approximately 15% for typical excess air percentages in steam boilers. (Broughton, 1994; "Economizers")

[[File:Economizer.png|thumb|border|center|middle|upright=4|link=|atl=|Economizer schematic ("Economizers") and economizer energy recovery correlations (Broughton Fig 2.2)]]

[[File:Condensate recycle.PNG|thumb|border|right|middle|200px|upright=4|link=|atl=|Condensate return fuel savings (Broughton Fig 2.4)]]

Additionally, there are still more opportunities for recovery in steam generation processes. One typical case is the use of waste heat to preheat air entering the boiler. Optimistically, this can result in a 5% improvement in heat recovery for the system. Another method is the recycle of steam condensate back to the boiler. This has two main benefits: it retains the heat still present in the condensed steam and it retains the pretreatment chemicals added to the steam. (Broughton, 1994)

Furthermore, there is opportunity for energy recovery in the expansion of compressed gas through a turbine to create electricity, a process that can be economically viable given sufficiently high flows or pressure. One of the primary application of this type of energy recovery is in the creation of medium pressure and low pressure steam. In most processes, all steam is generated as high pressure steam and can be expanded through a turbine to decrease its pressure. Such technology has also been used in processes to synthesize ammonia, perform air separations, and synthesize nitric acid. (Towler and Sinnott, 2012) Recently, however, there has been a particularly strong interest for energy recovery in the natural gas industry, when gas is decompressed from major pipelines to residential low-pressure pipelines. A 2001 study estimated that there is the potential to recover 21 TWh, representing 11% of natural gas transport energy, via gas expansion. (Lehman)

=Process water and boiler-feed water=
Process water is water that will be directly used in the process. Boiler-feed water (BFW) is used to produce steam. Both may need to be purified to prevent impurities from contaminating a process or from foul equipment. It can be used as a cooling stream when the temperature of the stream to be cooled is greater than ~300 ⁰F. Cost of BFW can be partially offset by the steam credit.

Process water that undergoes moderate pretreatment can cost ~ $0.75/1,000 gal.

Extensive treatment ~ $6.00/1,000 gal.

Sterilized for pharmaceutical processes ~ $550/1,000 gal. (Seider pg 608)

=Demineralized Water=
In demineralized water, minerals have been removed by ion exchange. In boiler feed water, this is to prevent salt deposition, corrosion, formation of foam, and sluicing. In process water, the ions may contaminate the process.

=Waste Treatment=
Most chemical processes will produce some sort of waste. Disposal occurs to the atmosphere (in the case of some gases), sewers, body of water, or a landfill. Waste may require some treatment before disposal to meet regulations. Depending on process economics, byproducts may be recovered and processed. (Seider 2009 pg 609)

==Wastewater Treatment==
The United States EPA regulates industrial wastewater disposal through the Clean Water Act, introduced in 1948 as the Federal Water Pollution Control Act and amended to its current form in 1972. The sweeping 1972 amendments allowed the EPA to prevent industries and persons from discharging contaminated water into fresh water sources and set water quality standards. (Summary of the Clean Water Act) In accordance with this law, process plants in the United States treat wastewater at on-site or near-site treatment centers before releasing it into the surrounding environment.

Wastewater effluent streams, along with water runoff from around the plant, are treated to control for pH, toxicity, suspended solids, and biological oxygen demand (for aquatic life protection) prior to discharge. Each of these controls is typically addressed with a separate method. Acidity and basicity is balanced through the addition of an acid or alkaline solution. Toxic wastewater may be treated with chemical processes or simply diluted to safe concentrations. Suspended solids can be removed via filtration and/or with clarifiers. Oxygen demand of wastewater can be mitigated using activated sludge treatment processes. Once the water quality complies with the EPA, and state-mandated, regulations, it can be safely released. More information on the large number of industry-specific guidelines for waste effluent can be found on the EPA website (http://www.epa.gov/eg/industrial-effluent-guidelines).

=Air-Pollution Management=
===Introduction===

In the United States air pollution is regulated in the Clean Air Act, and almost all pollutant emitting plants are regulated under this law. The types of plants that can release significant emissions include petroleum refineries, sulfur recovery plants, carbon-black plants, fuel conversion plants, chemical process plants, fossil fuel plants, and petroleum storage and transfer facilities. To receive permission to construct a plant must undergo a review to show that it will not cause a violation of the Ambient Air Quality Standards(Peters, 2003).

==Methods==

There are two major types of pollutants that are released into the air, particulates and and gaseous pollutants. Particulates can be removed with mechanical forces while gaseous pollutants typically need to removed by chemical or physical means (Peters, 2003).

Particulates and volatile pollutants that need to be removed before disposal may be present. Particle removal equipment includes: cyclones, wet scrubbers, electrostatic precipitators, and fabric-filter systems (Seider 2009) such as bag filters (. The two charts below are from ''Plant Design and Economics for Chemical Engineers'' and show the types of equipment, separation methods, and particle sizes in different pollutant separation technologies (Peters, 1991).

[[File:AirPolChart.png|thumb|border|center|middle|upright=4|link=|atl=|Different Methods of Particle Separation, the Particle Sizes they Can Remove, and the Technologies Used]]

[[File:AirPolTable.png|thumb|border|center|middle|upright=4|link=|atl=| This Chart Shows the Specifications and Limitations of Different Separations Technologies Including Particle Size, Efficiencies, and Temperatures(Peters, 1991)]]

Methods for removing inorganic and organic gaseous pollutants include: absorption, adsorption, condensation, and combustion (Seider 2009 pg 609). A list of typical of gases pollutants and their sources from ''Plant Design and Economics for Chemical Engineers'' is shown below (Peters, 1991).

[[File:GaseousPollutants.png|thumb|border|center|middle|upright=4|link=|atl=| Common Gaseous Pollutants and their Sources]]

Typically Gas-liquid absorption processes are done completed in a vertical, countercurrent, flow through packed, plate, or spray towers. These systems require good liquid-gas contact and proper equipment. These systems also often have significant energy consumption because of large pressure drops (Peters, 2003). For high volume systems absorption by scrubbing with water or another solvent is the most widely used method (Towler, 2012). Dry adsorbents can be used to remove the last races of gaseous pollutants. Adsorption typically requires blowers, condensers, separators, and controls. You also typically need two packed beds so that one can be used while the other is regenerated. Examples of adsorbents are molecular sieves and activated carbon. Incineration is typically used when there are gas streams that have no recovery value. This can be done with direct flame or catalytic oxidation. Catalytic oxidation usually has higher capital costs, but lower operating costs because it does not require fuel.
===Wet Scrubbers===
Wet scrubber use lime or limestone and water to remove SO2 and acid gases. The mixture can be injected into a scrubber or the gases can be bubbled through this mixture. This results in removal of 90-98% of SO2 and and acid gases (Clean Coal Technologies).

===Dry Scrubbers===
Dry scrubber blow powdered adsorbents into a vessel with gases and then after it has captured the SO2 and acid gases it is separated from gas using a fabric filter. These systems remove 90-93% of the contaminants (Clean Coal Technologies).

===Low-NOx Burners===
The purpose of low NOx burners is to decrease the amount of NOx created when the coal is burned. This is done by injecting coal and air in boilers. This can result in 40-50% NOx reduction. If air is injected into the area above the burner this can actually cause almost 70% NOx reduction (Clean Coal Technologies).

===Selective Non-Catalytic and Catalytic Reductions===
These systems inject ammonia into gases to remove NOx. The Catalytic reductions add a catalyst to the ammonia being injected to aid in the removal. Con-catalytic reductions result in about 35% removal, but adding a catalyst can increase that amount to about 90%. The catalytic-reduction can remove up to 80% of mercury as well (Clean Coal Technologies).

===Fabric Filters===
Also known as baghouses,these [[Separation processes#Filtration| filters]] remove particulates by passing air through filters. These can separate as much as 99.9% of particulate matter (Clean Coal Technologies).

===Electrostatic Precipitators===
Electrostatic Precicpitators remove particulate matter as gas passes through a device that has charged metal plates. The particles are then removed because of static electricity. These systems can remove between 99-99.9% of particulate matter (Clean Coal Technologies).

===Super-critical Boilers===
Super-Critical Boilers and Ultra-Supercritical Boilers operate at temperature and pressure higher than regular boilers. By operating at higher temperature these systems become more efficient. Super Critical Boilers typically have 10%-20% CO2 emissions than other similar sub-critical technologies. Ultra-Supercritical boilers can be as much as 30% more efficent than sub-critical technologies (Clean Coal Technologies).

===More Information===

For more information regarding some of these separations equipment see [[Solids-involved equipment]]

For more information Regarding Cyclones see [[Separation processes#Cyclones| Cyclones]], and for modeling cyclones see their[[Solids-involved equipment#HYSYS Simulation|HYSYS Simulation]]

You can find more information on [[Separation processes#Absorption| Absorption]] and [[Separation processes#Adsorption| Adsorption]] in [[Separation processes]]

===Outcomes===

====United States Common Pollutant Emissions====

The United States implemented the Clean Air Act in 1970 and since then emissions in the U.S. have been drastically reduced. As a matter of fact despite increases in the population in the last 35 years the amount of pollutants emitted have decreased by almost 70% (EPA). More importantly this demonstrates that reasonable efforts can be put towards environmental protection without causing too much harm to industry.

[[File:EPA.png|thumb|border|center|middle|upright=4|link=|atl=| GDP and Other Growth Factors Vs Common Pollutants Released]]

===China Clean Coal===

====Success Through 2005====

Coal is a very inexpensive and abundant source of energy and is abundant in China (Xu 2010). In China Coal the cause of 90% of SO2 emissions, 70% of dust emissions, and 67% of NOx emissions, and 70% of CO2 emissions. These numbers are staggering considering that Coal has never been more than 50 percent of China's Energy Supply (Xu 2010).

[[File:Coal.jpg|thumb|border|center|middle|upright=4|link=|atl=| Coal Consumption for Power and the Percentage of Coal Used (Xu 2010)]]

Despite the increasing coal consumption, high efficiency electric dust removal systems with efficiencies that are as high as 99.6% have greatly decreased soot emissions to 32% below 1980's levels as of 2005.

[[File:Soot.jpg|thumb|border|center|middle|upright=4|link=|atl=| SO2 and Soot Emissions from 1981 to 2005 (Xu 2010)]]

====Technologies and Life Cycle Comparison====
As of 2010 China was consuming 48.2% of coal globally. Four potential technologies integrated gasification combined cycle (IGCC), sub-critical coal power generation (Sub-C); super-critical coal power generation (Super-C) ultrasuper-critical coal power generation (USC). These technologies are compared mostly on the basis of net generating efficiency and efficiency. Net generating efficiency is the output of the plant divided by the total available energy in the fuel used (Liang, 2013).

where efficiency is define as: η=E/Etotal

and

Etotal=Emining+Etransportation+Egeneration

[[File:technologies.PNG|thumb|border|center|middle|upright=4|link=|atl=|A comparison of many of the parameters important to lfe-cycle analysis (Liang, 2013)]]

This demonstrates that USC has the best net generation efficiency, while also having the largest capacity for a single system.

[[File:Lifecycle2.PNG|thumb|border|center|middle|upright=4|link=|atl=| (Liang, 2013)]]

This breaks down the energy efficiency of each system and demonstrates that USC is the most efficient

[[File:Lifecycle.PNG|thumb|border|center|middle|upright=4|link=|atl=| (Liang, 2013)]]

This demonstrates that the capital cost per unit of energy production is very high for IGCC, but shows that Super-C and USC are very competitive.

[[File:Price.PNG|thumb|border|center|middle|upright=4|link=|atl=| Capital cost in $/kW of clean coal power-generation technologies (Liang, 2013)]]

Lastly IGCC and USC are lowest on different element, but since these all have different global warming potential it is difficult to tell which is the most efficient.

And the table below defines all the relevant terms.

[[File:Measures.PNG|thumb|border|center|middle|upright=4|link=|atl=| Environemntal measures and their explanations (Liang, 2013)]]

[[File:GWP.PNG|thumb|border|center|middle|upright=4|link=|atl=| Environemntal measures and their explanations (Liang, 2013)]]

All of the data suggests that USC is the highest in both energy efficiency and net generating efficiency while competitive in price and the lowest on emissions among the different boilers.

=Solid Waste=
U.S. federal regulations require that solid waste be classified as hazardous or nonhazardous. Conditions for a classification of hazardous include: ignitability, corrosivity, reactivity, toxicity, or posing a substantial threat to the surrounding environment and its inhabitants. Hazardous waste must be treated on- or near-site before being removed in containers. Non-hazardous waste may be landfilled or incinerated in some cases. A typical estimate of costs for waste disposal is $0.03/lb for nonhazardous solids and $0.10/lb for hazardous solids. (Seider 2009 pg 609)

=References=

*Biegler LT, Grossmann IE, Westerberg AW. ''Systematic Methods of Chemical Process Design'' Prentice-Hall: Upper Saddle River, 1997.

*Broughton, J. ''Process Utility Systems: Introduction to Design, Operation and Maintenance'' Institution of Chemical Engineers: Rugby, Warwickshire, UK, 1994.

*Clean Coal Technologies. America's Coalition for Clean Coal Electricity. http://www.americaspower.org/clean-coal-technologies-1663/

*Duke Energy Company (2013). How Do Coal Fired Plants Work? Charlotte: Duke Energy.

*"Economizers." Energy Solutions Center Inc.: Boiler Consortium. http://www.cleanboiler.org/Eff_Improve/Efficiency/Economizers.asp. Accessed February 21, 2016.

*Lehman B, Worrell E. ''Electricity Production from Natural Gas Pressure Recovery Using Expansion Turbines.'' Lawrence Berkeley National Laboratory; 2001.

*"Overview of the Clean Air Act and Air Pollution." Environmental Protection Agency. November 17, 2015. http://www.epa.gov/clean-air-act-overview. Accessed February 5, 2016.

*"Oxygen Control." Energy Solutions Center Inc.: Boiler Consortium. http://www.cleanboiler.org/Eff_Improve/Efficiency/Oxygen_Control.asp. Accessed February 21, 2016.

*Peters, Max S.; Timmerhaus, Klaus D.; West, Ronald E. (2003). " Plant Design and Economics for Chemical Engineers." McGraw Hill Higher Education.

*Ruina Xu Energy (2010). "Clean coal technology development in China." Elsevier.

*Seider, Seader, Lewin, Widagdo. (2009). ''Plant Design and Economics for Chemical Engineers, 5th Edition.'' Hoboken: Wiley.

*Seider, Seader, Lewin. (2008). ''Product and Process Design Principles, 2nd Edition.'' Hoboken: Wiley.

*Summary of the Clean Water Act. United States EPA website. http://www.epa.gov/laws-regulations/summary-clean-water-act

*Towler, G.P. and Sinnot, R. (2012). ''Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design.'' Elsevier.

*Turton R.; Bailie, R.C.; Whiting, W.B.; Shaeiwitz J.A.; Bhattacharyya D. (2012). ''Analysis, Synthesis, and Design of Chemical Processes.'' Upper Saddle River: Prentice Hall.

*G.D. Ulrich, A Guide to Chemical Engineering Process Design and Economics, Wiley: New York, 1984.

*Xiaoye Liang, Zhihua Wang, Zhijun Zhou, , Zhenyu Huang, Junhu Zhou, and Kefa Cen. (2013). "Up-to-date life cycle assessment and comparison study of clean coal power generation technologies in China" Elsevier.

Utility systems

2016-02-22T05:06:21Z

Bts511: /* Methods */

Authors: David Chen, [2014] Joshua Lee, [2015] Brett Sleyster, [2016] and Tom Aunins [2016] 

Stewards: David Chen, Jian Gong, and Fengqi You

Date Presented: January 13, 2014 /Date Revised: January 14, 2014

 


__TOC__

=Introduction=
Many chemical processes do not take place at ambient temperature or pressures. In order to reach these non-ambient conditions, utilities will have to be used to raise or lower temperatures and compress gases. (Towler, Towler/UOP) Utilities often contribute 5 to 10% of the price of a product, and may come from public or private utility companies or on-site plants. For purchased utilities, costs depend on consumption, while for company-owned utilities, there will be both capital and operating costs. They include things such as steam for heating, electricity, cooling water, refrigeration, fuels such as natural gas, wastewater treatment, waste disposal, and landfill. Steam is often the largest utility cost. Cogeneration unit can supply electricity accompanied with different steam pressures. (Seider 2009)

=Electricity=

Electricity is used to power many different kinds of equipment. It has many advantages: it is efficient (> 90%), reliable, available in a wide range of power, shaft speeds, designs, lifetimes, convenience, costs, and maintenance. It is generally used up to 200 hp, and sometimes over 10,000 Hp. In chemical process plants, the electricity demand is generally determined by the work or energy required for compression, pumping, air cooling, lights, and many other items. This electricity often times is purchased from local electricity providers, but many plants generate their own electricity via sophisticated processes.

Electricity is rarely used as a primary heat utility in large-scale chemical plants for a variety of reasons. The main disadvantages of using electricity as a heat utility are as follows (Towler 2012):

:*Heat from electricity is two to three times more expensive than heat from fuels. This is attributed to the lack of efficiency when creating heat from electricity.

:*Electrical heating units are expensive, require high maintenance, and must comply with strict safety regulations.

:*Electrical heating units are unsafe compared to steam heating units. In steam systems, the physically steam controls the temperature, whereas in electrical heating units temperature is controlled by temperature controllers, which can fail or burn out.

The use of electricity carries with it some hazards depending on the environment. Extra care must be taken when using electrically-powered equipment in areas which may have combustible fluids, vapors, or dust, and where liquids may be present. (Seider 2009 pg 606)

==Conventional Power Station==

In general, most electricity is generated from a conventional coal-fired process, whether it be on-site or purchased from a provider. Coal-fired processes have been used to create electricity throughout history, and technological advances have increased its efficiency and use worldwide. In a coal-fired steam station—much like a nuclear station—water is turned into steam, which in turn drives turbine generators to produce electricity. There are several variations on how to create energy from coal, but here are the basics of how a coal-fired process works:

[[File:coalfired.png|thumb|border|center|middle|upright=4|link=|atl=|General Coal-Fired Process Diagram (Duke Energy Company)]]

*Heat is created:
::Before the coal is burned, it is pulverized to the fineness of talcum powder. It is then mixed with hot air and blown into the firebox of the boiler. Burning in suspension, the coal/air mixture provides the most complete combustion and maximum heat possible.

*Water turns to steam:
::Highly purified water, pumped through pipes inside the boiler, is turned into steam by the heat. The steam reaches temperatures of up to 1,000 degrees Fahrenheit and pressures up to 3,500 pounds per square inch, and is piped to the turbine.

*Steam turns the turbine:
::The enormous pressure of the steam pushing against a series of giant turbine blades turns the turbine shaft. The turbine shaft is connected to the shaft of the generator, where magnets spin within wire coils to produce electricity.

*Steam is converted back to water:
::After doing its work in the turbine, the steam is drawn into a condenser, a large chamber in the basement of the power plant. In this important step, millions of gallons of cool water from a nearby source (such as a river or lake) are pumped through a network of tubes running through the condenser. The cool water in the tubes converts the steam back into water that can be used over and over again in the plant.

*Repeat:
::The cooling water is returned to its source without any contamination, and the steam water is returned to the boiler to repeat the cycle.

===Advantages & Disadvantages of Coal-Fired Energy Production===

[[File:coalfiredtable.png|thumb|border|center|middle|upright=4|link=|atl=|Advantages and Disadvantages of Coal-Fired Processes (Seider 2008)]]

==Gas-Turbine Cogeneration Process==

When generating energy on-site, many plants use a gas-turbine cogeneration process. The thermal efficiency of a gas-turbine process is in the range of 70-80% while conventional power stations, such as coal-fired processes, have a 30-40% efficiency. The lower efficiency in more conventional power stations is attributed to wasted heat in the exhaust steam in the condenser. One example of a gas-turbine process is outlined in the following figure. Figure 3.1 is a gas-turbine cogeneration process with a heat recovery steam generator (waste-heat) boiler.

[[File:cogeneration.png|thumb|border|center|middle|upright=4|link=|atl=|Gas-Turbine Cogeneration with a heat recovery steam generator boiler (Towler 2012 Fig 3.1)]]

Overall, the process illustrated is not much different from a coal-fired process. The main differences are that the cogeneration process creates both electricity and a heat utility, and cogeneration processes use natural gas instead of coal. Many of the advantages and disadvantages are similar to those of the coal-fired process, but the cogeneration has a much higher efficiency, creates heat to be used in another process, and uses a more volatile and expensive fuel. The main advantage of cogeneration over coal-fired energy production is that heat is not wasted. In coal-fired processes, heat is released and wasted during electricity generation. Cogeneration captures some, if not all of the byproduct for heat, which is an extremely useful utility that will be discussed in subsequent sections. In summary, the cogeneration plant is superior to the coal-fired process because of its higher efficiency and ability to create a useful heat utility.

Obviously any engineer would design the cogeneration plant to meet at least the energy requirement necessary for plant operation, but cogeneration plants often times are designed to exceed the plant electricity requirement to drive another source of capital. Many describe this scenario as a "make or buy" scenario (Towler 2012). This scenario provides chemical producers leverage when negotiating contracts with outsourced electricity providers and this allows plants to purchase electricity at a wholesale price. This is a huge advantage of considering on-site electricity production because electricity is needed in relatively high quantities for all chemical plants. Being able to minimize electricity costs, or even profit off of electricity production is a huge economical consideration that all plants employ.

=Process Heating=
The key objective of this section is to discuss how processes are heated. Heating utilities are necessary for proper usage of distillers, reactors, condensers, and several other integral types of equipment. More specifically, steam, fired heat, and hot oil/specialized heat transfer fluids will be discussed in the following subsections.

==Steam==
Steam is the most commonly used heat utility used in chemical plants, and as a result understanding how it is used is essential in the study of Utility systems. Steam is used both as a process fluid (feedstock, diluent to absorb heat of reaction, heating agent, and stripping agent in absorbers and adsorbers ) and utility. It can be used to drive pumps and compressors, ejectors (for producing a vacuum), and heat exchangers. As one can clearly see, steam is a versatile, and useful utility.

Here are a few advantages of using steam as opposed to other methods of process heating (Towler 2012):

:*By controlling the pressure of the steam, one can control the temperature at which the heat is released. Having a strong control over the temperature is essential in several processes. =
:*Steam is an efficient heat source because the heat of condensation of steam is very high. Meaning that there is is high output per mass of utility at a constant temperature.
:*Heat exchangers that use steam are relatively cheap because condensing steam has a high heat transfer coefficient.
:*Steam is nonflammable, nontoxic, and inert to several process fluids.

Chemical plants generally have a network of pipelines exclusively for providing steam. These networks generally have steam at a low pressure, a medium pressure, and a high pressure. The image below illustrates a basic steam system.

[[File:Steam.png|thumb|border|center|middle|upright=4|link=|atl=|(Towler 2012 Fig 3.2)]]

In the diagram above, boiler feed water at a high pressure is preheated and fed to other boilers. These other boilers superheat the steam to create a high pressure and high temperature steam stream. The steam is superheated past the dew point to account for heat loss in the pipelines. A portion of the high pressure steam is used for process heating in areas of the plant that require high temperatures. The rest of the high pressure steam is turned into medium pressure steam by valves and steam turbines. The medium pressure steam is then used to heat medium temperature processes and to form low pressure steam. The low pressure steam can be used to heat low pressure processes and it can be expanded in condensing turbines to create shaft work and energy. In summary, steam can be used for an innumerable amount of action items in a plant. High pressure, medium pressure, and low pressure steam can all be used as a heat source. Low pressure steam has utilities in creating electricity and it also has several other uses.

==Fired Heat==

In many cases, processes in a plant require a heat source stronger than high pressure and temperature steam. That is when fired heat is used, which is generally at temperatures above 523K. Streams can be heated directly in the furnace tubes or via a hot oil circuit or heat transfer fluid, which will be discussed in detail in the next section. Most fired heaters use natural gas as fuel because it burns cleaner than fuel oil. A cleaner burning fuel is always advantageous due to environmental and safety concerns. Furthermore, natural gases usually result in less wear and tear in burners and fuel lines.

Depending on the application of the fired-heater, different design specifications can be implemented to make the fired-heater as efficient as possible. The basic construction of a fired heater starts with a cylindrical chamber that is lined with with refractory bricks. The stream that is to be heated flows through tubes inside of the furnace. These tubes can be arranged in several different arrangements such as, around the walls of the furnace, or in horizontal or vertical banks (Towler 2012). The figure below illustrates the basic construction of the fired-heater and varying tube arrangements.

[[File:firedheater.png|thumb|border|center|middle|upright=4|link=|atl=|Basic Fired-Heater Diagrams (Towler 2012 Figure 19.66)]]

Fuel is burned to heat the entire furnace, and the heat transfer occurs from the combustion gases inside of the furnace across the tubes that are filled with our desired stream. The heat transfer between the tubes and the furnace is accomplished primarily via radiation. Modern designs take advantage of convective heat transfer by adding a smaller section on top of the fired-heater where the combustion gases flow over banks of tubes as seen in (c) in the figure above. Heat transfer can be further improved via convection by adding fins or pins in the combustion section.

The cost of fired heat can be calculated by the cost of fuel fired. Natural gas and heating oil are traded as commodities and prices can be found at many online trading sites or business news sites (i.e., www.cnn.money.com). Past historic prices for forecasting can be found in the Oil and Gas Journal or from the U.S. Energy Information Adminstration (www.eia.gov).

===Fuels===
Fuel is burned in utility facilities such as boilers, electricity generation, and cogeneration, and can be in solid, liquid, or gas form. It can also be burned to provide heating for a process or stream or to drive pumps and compressors. The fuel is usually burned with excess air to ensure complete combustion.

A way of quantifying the amount of heat generated is by using the heating values. Higher heating value (HHV) and the lower heating value (LHV) are used. The heating is the total heat evolved by complete combustion of a fuel with dry air with both at 60 ⁰F and the flue gas after combustion brought back down to ⁰F. If the water vapor in the flue gas is not condensed, we obtain the LHV. If the water vapor is condensed, the value of heat evolved is a bit higher, and this is the HHV.
Heating values for solids and liquids are usually on a per-mass basis, and gases on a per-volume basis.
A specified amount of heating cannot be met with the amount of fuel calculated using only the HHV. There will be heat losses, the flue gas temperature will be greater than 60 ⁰F, and water in the flue gas will typically be vapor. (Seider 608)

==Hot Oil/Specialized Heat Transfer Fluids==

Specialized heat transfer fluids and hot oil circuits are used as heat sources when steam and fired heat is not appropriate. Specialized heat transfer fluids and hot oil circuits are extremely versatile in that they can be used in the temperature range of 323K to 673K. This range however is quite variable. For hot oils, the upper temperature limit is gauged based off of the thermal decomposition of the oil and coking/fouling of heat exchanger tubes.

Hot oil circuit systems are most commonly used when the plant has several small temperature heating requirements because it is more economically sound. Rather than having several fired heaters heat each small temperature requirement, it is much more economical to have one fired heater heat the hot oil and circulate that oil through each of the process to meet all of the heat needs. Hot oil systems are also generally favored over high pressure steam in processes that involve high pressure differentials between the process stream and high pressure steam. Hot oil systems are favored in this scenario because of safety concerns. If the steam were to leak, the pressure drop could cause serious safety issues.

Mineral oils are the most commonly used heat transfer fluids, and one prominent example is Dowtherm A. Dowtherm A is a combination of 26.5 wt% diphenyl in diphenyl oxide (Towler 2012) and is extremely thermally stable. These mineral oil systems generally require high flow rates.

=Process Cooling=

==Cooling Water==
Cooling water is used to cool and/or condense streams. Cooling water is usually circulated between process heat exchangers and a cooling tower. Water is cooled during downward motion by contact with air blown upwards, which can bring the water temperature to come within ~ 5 ⁰F of air’s wet-bulb temperature.Approximately 80% of the temperature reduction is due to evaporation of the cooling water and heat transfer to the surrounding air. Water can also be cooled in spray ponds and cooling ponds. Both work by providing high area for water to exchange heat with air.
Water in cooling towers is lost through drift and blowdown, and makeup is usually 1.5 to 3% of the circulating rate.
If a large natural body of water is nearby, it can be used as a source of cooling water and discharged downstream. This water is usually filtered to remove salts and impurities that may lead to fouling, but it is not treated.

==Refrigeration==
Cooling water can usually be used to cool a stream to ~ 100 ⁰F. Air can only cool to ~ 120 ⁰F. Air may be used in places where water is scarce or more costly to transport. To cool or condense streams below 100 ⁰F, a refrigerant is typically used. Chilled brine can also be used, but is less common.

Until 1995, CFC Freon R-12 (dichlorodifuloromethane) and HCFC Freon R-22 (chlorodifuloromethane) were commonly used refrigerants. However, the chlorine atom in the molecules caused the depletion of the ozone layer. The U.S. Clean Air Act Amendments of 1990 went into effect in 1995, and the production of these refrigerants has since stopped or been greatly reduced.

Cost estimates are based on ton-day of refrigeration, where a ton is the heat that needs to be removed to freeze 1 ton per day of water at 32 ⁰F. Substitutes have since been developed. R-134a is often used in place of R-12. According to the EPA, R-134a is not combustible at ambient conditions, and is essentially non-toxic under 400 ppm, and is not ozone-depleting. (Seider pg 607)

=Energy Efficiency=
One of the chief concerns in selecting and designing process utility systems for heating and cooling is how to achieve the most energy efficient design. There are countless means by which plants lose energy, two of the foremost being through the mixing of different temperature or pressure streams and through the disposal of warmed cooling water. (Seider, 2009) Proper utilities design can help mitigate each of these losses as well as many others. The energy efficiency of a plant will depend primarily on the heating and cooling methods that are being used and the overall system design itself. These two parameters are important in determining how well energy is transferred to the desired media as well as how well that energy is recovered and recycled.

==Hot Utility Efficiency==
As mentioned above, the most commonly used utilities for process heating in large scale processes are steam, fired heat, and hot oil heaters. Of these, steam is the most commonly used. Electricity, while efficient at creating power, is not a viable source of heat in large industrial processes. Common ranges of heating efficiency for these three methods are displayed in Table 1. (Towler and Sinnott, 2012; Broughton, 1994)

{| class="wikitable"
|+Table 1: Process Heating Efficiencies
|-
! Process Heating Method
! Typical Efficiency
|-
| Steam (via coal boiler)
| 72%
|-
| Steam (via gas boiler)
| 73%
|-
| Steam (via oil boiler)
| 75%
|-
| Fired Heat w/ Convective Section
| 85%
|-
| Fired Heat w/o Convective Section
| 60%
|-
| Hot Oil Heaters/Vaporizers
| 80-85%
|}

===Steam===
As steam is so popular for heating purposes, it is useful to analyze the numerous ways in which losses can occur in steam systems. There are five primary sources of inefficiency and heat loss in the generation and distribution of steam throughout a process plant:

*The heat content of boiler exhaust gas
*Incomplete combustion of boiler fuel
*Radiant losses from the boiler exterior
*Blowdown losses
*Distribution losses (pipe transport, steam traps, etc.)

The first four of these losses take place at the boiler and contribute to the heating efficiencies seen in Table 1 for steam created with coal, gas, and oil. Several methods can be used to minimize these losses, one of the foremost being the control of air-to-fuel ratio in the boiler. This ratio must be managed by weighing losses due to uncombusted fuel against losses due to the heat content of excess exhaust gas. (Broughton, 1994)

[[File:Air-fuel ratio.PNG|thumb|border|center|middle|upright=4|link=|atl=|Air-to-fuel ratio relationship with heat losses (Broughton Fig 2.3)]]

The air-to-fuel ratio can be optimized using a feedback process controller. The control system will analyze the oxygen content of exhaust air and adjust the incoming air flow rate to achieve a set percentage of excess air. While desired excess oxygen will vary depending on the type of fuel, it is consistently seen that in the minimum loss-range a 1% increase in excess air will result in a 1% decrease in efficiency. ("Oxygen Control")

In addition to air-to-fuel ratio management, steam heat losses at the boiler can be mitigated via energy recovery, which is discussed in further detail [[Utility_systems#Energy_Recovery|below]].

Outside the boiler, losses in distribution of steam throughout facilities can decrease energy efficiency by a significant amount. This can cause up to a 60% increase in losses, but typically results in an overall steam efficiency of 50-55% (down from approximately 75% at the boiler). (Broughton, 1994) There are several ways that this issue can be addressed. First and foremost are steam metering systems, which can be used to monitor heat supply to separate sections of the process facility, check efficiency of fuel use, and determine which processes in a given plant have the highest steam requirements. Another method of minimizing transport losses is to decentralize steam generation systems. It can be advantageous to have numerous smaller boilers rather than a single centralized boiler so that steam does not have to travel as long of distances to reach its destination.

==Cold Utility Efficiency==
Efficiency in cooling processes depends more on the method used, and by extension the amount of coolant needed. Water and air utility efficiencies depend primarily on the fluid flow required to maintain the system at a desired temperature, while powered refrigeration utilities (for colder processes) have efficiencies at approximately 60%--but ranging up to 90%--of Carnot cycle efficiency, a metric of ideal refrigeration efficiency. (Towler and Sinnott, 2012) Cooling systems represent by definition a loss of energy from the main process to the utility stream, and as such it is often useful to find other uses for the heated media before discharge.

==Energy Recovery==
Recovery and recycle of energy is perhaps the most important aspect of creating an energy efficient plant design, and it is important for process engineers to fully consider possibilities for heat recovery in order to aid in economic viability.

===Process Heat Exchange===
Heat exchanger networks are a very common energy recovery method in industrial processes. These networks most frequently allow energy from heated product streams to be transferred to feed streams that must be brought up to process temperature. (Biegler, 1997) More information on the function and design of heat exchanger networks can be found on the [[Heat_exchanger#Heat_Exchanger_Networks|heat exchanger wiki page]]. The following are several examples of energy recovery via heat exchange that are used in industrial processes.

In distillation columns the bottoms and distillate effluents have the potential for energy exchange. Though the condenser at the top of the column cannot supply its waste heat to the reboiler due to their respective temperatures, the effluent streams can supply heat to the feed via a feed-effluent exchanger. This reduces the utility requirements to raise the feed to column temperature. (Biegler, 1997)

[[File:Feed_sterilization.jpg|thumb|border|right|middle|upright=4|link=|atl=|Feed sterilization schematic. (Towler and Sinnott Fig. 3.30)]]

Feed sterilization, commonly used in the food industry, is a common application for heat recovery through process stream heat exchange. In this application, the feed must be heated for a certain amount of time to kill any biological contaminants, after which it can be used to heat the new raw feed for sterilization. This reduces energy demands on the steam heater and thus reduces cost. (Towler and Sinnott, 2012)

In multi-vessel batch processes it can be advantageous to exchange heat as the process fluid is being transferred between vessels. Like the previous examples, this reduces the utility needed to bring the colder feed up to process temperature, thus reducing costs. (Towler and Sinnott, 2012)

[[File:Batch_heat_exchange.jpg|thumb|border|center|middle|upright=4|link=|atl=|Multi-vessel batch heat exchange schematic (Towler and Sinnott Fig. 3.31)]]

===Utility Regeneration===
[[File:Waste_heat_boiler.jpg|thumb|border|right|middle|200px|upright=4|link=|atl=|Industrial modular HRSG]]
When recovery of waste heat via transfer to other process streams is inconvenient or impossible, energy efficiency can still be improved through the regeneration of utilities. This is commonly done through the regeneration of steam by removing heat from exiting streams or from highly exothermic reactions. Waste heat in exiting streams can be removed via heat recovery steam generators (HRSGs), and is most often used on exiting gas streams. Heat recovery from reactions is a viable option when the reactor temperature will be at 150 C or above, as this will create steam at high enough pressure to be used in other processes. (Towler and Sinnott, 2012)

In the case of steam, waste heat and water treatment losses can be recovered from the utility generation process itself. One of the most common ways to do this is with an economizer. As seen in the schematic below, economizers have heat exchangers fit to the exhaust gas flow in order to transfer waste heat from these gases to the incoming boiler feed water. This can result in fuel energy savings of approximately 15% for typical excess air percentages in steam boilers. (Broughton, 1994; "Economizers")

[[File:Economizer.png|thumb|border|center|middle|upright=4|link=|atl=|Economizer schematic ("Economizers") and economizer energy recovery correlations (Broughton Fig 2.2)]]

[[File:Condensate recycle.PNG|thumb|border|right|middle|200px|upright=4|link=|atl=|Condensate return fuel savings (Broughton Fig 2.4)]]

Additionally, there are still more opportunities for recovery in steam generation processes. One typical case is the use of waste heat to preheat air entering the boiler. Optimistically, this can result in a 5% improvement in heat recovery for the system. Another method is the recycle of steam condensate back to the boiler. This has two main benefits: it retains the heat still present in the condensed steam and it retains the pretreatment chemicals added to the steam. (Broughton, 1994)

Furthermore, there is opportunity for energy recovery in the expansion of compressed gas through a turbine to create electricity, a process that can be economically viable given sufficiently high flows or pressure. One of the primary application of this type of energy recovery is in the creation of medium pressure and low pressure steam. In most processes, all steam is generated as high pressure steam and can be expanded through a turbine to decrease its pressure. Such technology has also been used in processes to synthesize ammonia, perform air separations, and synthesize nitric acid. (Towler and Sinnott, 2012) Recently, however, there has been a particularly strong interest for energy recovery in the natural gas industry, when gas is decompressed from major pipelines to residential low-pressure pipelines. A 2001 study estimated that there is the potential to recover 21 TWh, representing 11% of natural gas transport energy, via gas expansion. (Lehman)

=Process water and boiler-feed water=
Process water is water that will be directly used in the process. Boiler-feed water (BFW) is used to produce steam. Both may need to be purified to prevent impurities from contaminating a process or from foul equipment. It can be used as a cooling stream when the temperature of the stream to be cooled is greater than ~300 ⁰F. Cost of BFW can be partially offset by the steam credit.

Process water that undergoes moderate pretreatment can cost ~ $0.75/1,000 gal.

Extensive treatment ~ $6.00/1,000 gal.

Sterilized for pharmaceutical processes ~ $550/1,000 gal. (Seider pg 608)

=Demineralized Water=
In demineralized water, minerals have been removed by ion exchange. In boiler feed water, this is to prevent salt deposition, corrosion, formation of foam, and sluicing. In process water, the ions may contaminate the process.

=Waste Treatment=
Most chemical processes will produce some sort of waste. Disposal occurs to the atmosphere (in the case of some gases), sewers, body of water, or a landfill. Waste may require some treatment before disposal to meet regulations. Depending on process economics, byproducts may be recovered and processed. (Seider 2009 pg 609)

==Wastewater Treatment==
The United States EPA regulates industrial wastewater disposal through the Clean Water Act, introduced in 1948 as the Federal Water Pollution Control Act and amended to its current form in 1972. The sweeping 1972 amendments allowed the EPA to prevent industries and persons from discharging contaminated water into fresh water sources and set water quality standards. (Summary of the Clean Water Act) In accordance with this law, process plants in the United States treat wastewater at on-site or near-site treatment centers before releasing it into the surrounding environment.

Wastewater effluent streams, along with water runoff from around the plant, are treated to control for pH, toxicity, suspended solids, and biological oxygen demand (for aquatic life protection) prior to discharge. Each of these controls is typically addressed with a separate method. Acidity and basicity is balanced through the addition of an acid or alkaline solution. Toxic wastewater may be treated with chemical processes or simply diluted to safe concentrations. Suspended solids can be removed via filtration and/or with clarifiers. Oxygen demand of wastewater can be mitigated using activated sludge treatment processes. Once the water quality complies with the EPA, and state-mandated, regulations, it can be safely released. More information on the large number of industry-specific guidelines for waste effluent can be found on the EPA website (http://www.epa.gov/eg/industrial-effluent-guidelines).

=Air-Pollution Management=
===Introduction===

In the United States air pollution is regulated in the Clean Air Act, and almost all pollutant emitting plants are regulated under this law. The types of plants that can release significant emissions include petroleum refineries, sulfur recovery plants, carbon-black plants, fuel conversion plants, chemical process plants, fossil fuel plants, and petroleum storage and transfer facilities. To receive permission to construct a plant must undergo a review to show that it will not cause a violation of the Ambient Air Quality Standards(Peters, 2003).

==Methods==

There are two major types of pollutants that are released into the air, particulates and and gaseous pollutants. Particulates can be removed with mechanical forces while gaseous pollutants typically need to removed by chemical or physical means (Peters, 2003).

Particulates and volatile pollutants that need to be removed before disposal may be present. Particle removal equipment includes: cyclones, wet scrubbers, electrostatic precipitators, and fabric-filter systems (Seider 2009) such as bag filters (. The two charts below are from ''Plant Design and Economics for Chemical Engineers'' and show the types of equipment, separation methods, and particle sizes in different pollutant separation technologies (Peters, 1991).

[[File:AirPolChart.png|thumb|border|center|middle|upright=4|link=|atl=|Different Methods of Particle Separation, the Particle Sizes they Can Remove, and the Technologies Used]]

[[File:AirPolTable.png|thumb|border|center|middle|upright=4|link=|atl=| This Chart Shows the Specifications and Limitations of Different Separations Technologies Including Particle Size, Efficiencies, and Temperatures(Peters, 1991)]]

Methods for removing inorganic and organic gaseous pollutants include: absorption, adsorption, condensation, and combustion (Seider 2009 pg 609). A list of typical of gases pollutants and their sources from ''Plant Design and Economics for Chemical Engineers'' is shown below (Peters, 1991).

[[File:GaseousPollutants.png|thumb|border|center|middle|upright=4|link=|atl=| Common Gaseous Pollutants and their Sources]]

Typically Gas-liquid absorption processes are done completed in a vertical, countercurrent, flow through packed, plate, or spray towers. These systems require good liquid-gas contact and proper equipment. These systems also often have significant energy consumption because of large pressure drops (Peters, 2003). For high volume systems absorption by scrubbing with water or another solvent is the most widely used method (Towler, 2012). Dry adsorbents can be used to remove the last races of gaseous pollutants. Adsorption typically requires blowers, condensers, separators, and controls. You also typically need two packed beds so that one can be used while the other is regenerated. Examples of adsorbents are molecular sieves and activated carbon. Incineration is typically used when there are gas streams that have no recovery value. This can be done with direct flame or catalytic oxidation. Catalytic oxidation usually has higher capital costs, but lower operating costs because it does not require fuel.
===Wet Scrubbers===
Wet scrubber use lime or limestone and water to remove SO2 and acid gases. The mixture can be injected into a scrubber or the gases can be bubbled through this mixture. This results in removal of 90-98% of SO2 and and acid gases (Clean Coal Technologies).

===Dry Scrubbers===
Dry scrubber blow powdered adsorbents into a vessel with gases and then after it has captured the SO2 and acid gases it is separated from gas using a fabric filter. These systems remove 90-93% of the contaminants (Clean Coal Technologies).

===Low-NOx Burners===
The purpose of low NOx burners is to decrease the amount of NOx created when the coal is burned. This is done by injecting coal and air in boilers. This can result in 40-50% NOx reduction. If air is injected into the area above the burner this can actually cause almost 70% NOx reduction (Clean Coal Technologies).

===Selective Non-Catalytic and Catalytic Reductions===
These systems inject ammonia into gases to remove NOx. The Catalytic reductions add a catalyst to the ammonia being injected to aid in the removal. Con-catalytic reductions result in about 35% removal, but adding a catalyst can increase that amount to about 90%. The catalytic-reduction can remove up to 80% of mercury as well (Clean Coal Technologies).

===Fabric Filters===
Also known as baghouses,these [[Separation processes#Filtration| filters]] remove particulates by passing air through filters. These can separate as much as 99.9% of particulate matter (Clean Coal Technologies).

===Electrostatic Precipitators===
Electrostatic Precicpitators remove particulate matter as gas passes through a device that has charged metal plates. The particles are then removed because of static electricity. These systems can remove between 99-99.9% of particulate matter (Clean Coal Technologies).

===Super-critical Boilers===
Super-Critical Boilers and Ultra-Supercritical Boilers operate at temperature and pressure higher than regular boilers. By operating at higher temperature these systems become more efficient. Super Critical Boilers typically have 10%-20% CO2 emissions than other similar sub-critical technologies. Ultra-Supercritical boilers can be as much as 30% more efficent than sub-critical technologies (Clean Coal Technologies).

===More Information===

For more information regarding some of these separations equipment see [[Solids-involved equipment]]

For more information Regarding Cyclones see [[Separation processes#Cyclones| Cyclones]], and for modeling cyclones see their[[Solids-involved equipment#HYSYS Simulation|HYSYS Simulation]]

You can find more information on [[Separation processes#Absorption| Absorption]] and [[Separation processes#Adsorption| Adsorption]] in [[Separation processes]]

===Outcomes===

====United States Common Pollutant Emissions====

The United States implemented the Clean Air Act in 1970 and since then emissions in the U.S. have been drastically reduced. As a matter of fact despite increases in the population in the last 35 years the amount of pollutants emitted have decreased by almost 70% (EPA). More importantly this demonstrates that reasonable efforts can be put towards environmental protection without causing too much harm to industry.

[[File:EPA.png|thumb|border|center|middle|upright=4|link=|atl=| GDP and Other Growth Factors Vs Common Pollutants Released]]

===China Clean Coal===

====Success Through 2005====

Coal is a very inexpensive and abundant source of energy and is abundant in China (Xu 2010). In China Coal the cause of 90% of SO2 emissions, 70% of dust emissions, and 67% of NOx emissions, and 70% of CO2 emissions. These numbers are staggering considering that Coal has never been more than 50 percent of China's Energy Supply (Xu 2010).

[[File:Coal.jpg|thumb|border|center|middle|upright=4|link=|atl=| Coal Consumption for Power and the Percentage of Coal Used (Xu 2010)]]

Despite the increasing coal consumption, high efficiency electric dust removal systems with efficiencies that are as high as 99.6% have greatly decreased soot emissions to 32% below 1980's levels as of 2005.

[[File:Soot.jpg|thumb|border|center|middle|upright=4|link=|atl=| SO2 and Soot Emissions from 1981 to 2005 (Xu 2010)]]

====Technologies and Life Cycle Comparison====
As of 2010 China was consuming 48.2% of coal globally. Four potential technologies integrated gasification combined cycle (IGCC), sub-critical coal power generation (Sub-C); super-critical coal power generation (Super-C) ultrasuper-critical coal power generation (USC). These technologies are compared mostly on the basis of net generating efficiency and efficiency. Net generating efficiency is the output of the plant divided by the total available energy in the fuel used (Liang, 2013).

where efficiency is define as: η=E/Etotal

and

Etotal=Emining+Etransportation+Egeneration

[[File:technologies.PNG|thumb|border|center|middle|upright=4|link=|atl=|A comparison of many of the parameters important to lfe-cycle analysis (Liang, 2013)]]

This demonstrates that USC has the best net generation efficiency, while also having the largest capacity for a single system.

[[File:Lifecycle2.PNG|thumb|border|center|middle|upright=4|link=|atl=| (Liang, 2013)]]

This breaks down the energy efficiency of each system and demonstrates that USC is the most efficient

[[File:Lifecycle.PNG|thumb|border|center|middle|upright=4|link=|atl=| (Liang, 2013)]]

This demonstrates that the capital cost per unit of energy production is very high for IGCC, but shows that Super-C and USC are very competitive.

[[File:Price.PNG|thumb|border|center|middle|upright=4|link=|atl=| Capital cost in $/kW of clean coal power-generation technologies (Liang, 2013)]]

Lastly IGCC and USC are lowest on different element, but since these all have different global warming potential it is difficult to tell which is the most efficient.

And the table below defines all the relevant terms.

[[File:GWP.PNG|thumb|border|center|middle|upright=4|link=|atl=| Environemntal measures and their explanations (Liang, 2013)]]

All of the data suggests that USC is the highest in both energy efficiency and net generating efficiency while competitive in price and the lowest on emissions among the different boilers.

=Solid Waste=
U.S. federal regulations require that solid waste be classified as hazardous or nonhazardous. Conditions for a classification of hazardous include: ignitability, corrosivity, reactivity, toxicity, or posing a substantial threat to the surrounding environment and its inhabitants. Hazardous waste must be treated on- or near-site before being removed in containers. Non-hazardous waste may be landfilled or incinerated in some cases. A typical estimate of costs for waste disposal is $0.03/lb for nonhazardous solids and $0.10/lb for hazardous solids. (Seider 2009 pg 609)

=References=

*Biegler LT, Grossmann IE, Westerberg AW. ''Systematic Methods of Chemical Process Design'' Prentice-Hall: Upper Saddle River, 1997.

*Broughton, J. ''Process Utility Systems: Introduction to Design, Operation and Maintenance'' Institution of Chemical Engineers: Rugby, Warwickshire, UK, 1994.

*Clean Coal Technologies. America's Coalition for Clean Coal Electricity. http://www.americaspower.org/clean-coal-technologies-1663/

*Duke Energy Company (2013). How Do Coal Fired Plants Work? Charlotte: Duke Energy.

*"Economizers." Energy Solutions Center Inc.: Boiler Consortium. http://www.cleanboiler.org/Eff_Improve/Efficiency/Economizers.asp. Accessed February 21, 2016.

*Lehman B, Worrell E. ''Electricity Production from Natural Gas Pressure Recovery Using Expansion Turbines.'' Lawrence Berkeley National Laboratory; 2001.

*"Overview of the Clean Air Act and Air Pollution." Environmental Protection Agency. November 17, 2015. http://www.epa.gov/clean-air-act-overview. Accessed February 5, 2016.

*"Oxygen Control." Energy Solutions Center Inc.: Boiler Consortium. http://www.cleanboiler.org/Eff_Improve/Efficiency/Oxygen_Control.asp. Accessed February 21, 2016.

*Peters, Max S.; Timmerhaus, Klaus D.; West, Ronald E. (2003). " Plant Design and Economics for Chemical Engineers." McGraw Hill Higher Education.

*Ruina Xu Energy (2010). "Clean coal technology development in China." Elsevier.

*Seider, Seader, Lewin, Widagdo. (2009). ''Plant Design and Economics for Chemical Engineers, 5th Edition.'' Hoboken: Wiley.

*Seider, Seader, Lewin. (2008). ''Product and Process Design Principles, 2nd Edition.'' Hoboken: Wiley.

*Summary of the Clean Water Act. United States EPA website. http://www.epa.gov/laws-regulations/summary-clean-water-act

*Towler, G.P. and Sinnot, R. (2012). ''Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design.'' Elsevier.

*Turton R.; Bailie, R.C.; Whiting, W.B.; Shaeiwitz J.A.; Bhattacharyya D. (2012). ''Analysis, Synthesis, and Design of Chemical Processes.'' Upper Saddle River: Prentice Hall.

*G.D. Ulrich, A Guide to Chemical Engineering Process Design and Economics, Wiley: New York, 1984.

*Xiaoye Liang, Zhihua Wang, Zhijun Zhou, , Zhenyu Huang, Junhu Zhou, and Kefa Cen. (2013). "Up-to-date life cycle assessment and comparison study of clean coal power generation technologies in China" Elsevier.

Utility systems

2016-02-22T05:02:02Z

Bts511: /* Solid Waste */

Authors: David Chen, [2014] Joshua Lee, [2015] Brett Sleyster, [2016] and Tom Aunins [2016] 

Stewards: David Chen, Jian Gong, and Fengqi You

Date Presented: January 13, 2014 /Date Revised: January 14, 2014

 


__TOC__

=Introduction=
Many chemical processes do not take place at ambient temperature or pressures. In order to reach these non-ambient conditions, utilities will have to be used to raise or lower temperatures and compress gases. (Towler, Towler/UOP) Utilities often contribute 5 to 10% of the price of a product, and may come from public or private utility companies or on-site plants. For purchased utilities, costs depend on consumption, while for company-owned utilities, there will be both capital and operating costs. They include things such as steam for heating, electricity, cooling water, refrigeration, fuels such as natural gas, wastewater treatment, waste disposal, and landfill. Steam is often the largest utility cost. Cogeneration unit can supply electricity accompanied with different steam pressures. (Seider 2009)

=Electricity=

Electricity is used to power many different kinds of equipment. It has many advantages: it is efficient (> 90%), reliable, available in a wide range of power, shaft speeds, designs, lifetimes, convenience, costs, and maintenance. It is generally used up to 200 hp, and sometimes over 10,000 Hp. In chemical process plants, the electricity demand is generally determined by the work or energy required for compression, pumping, air cooling, lights, and many other items. This electricity often times is purchased from local electricity providers, but many plants generate their own electricity via sophisticated processes.

Electricity is rarely used as a primary heat utility in large-scale chemical plants for a variety of reasons. The main disadvantages of using electricity as a heat utility are as follows (Towler 2012):

:*Heat from electricity is two to three times more expensive than heat from fuels. This is attributed to the lack of efficiency when creating heat from electricity.

:*Electrical heating units are expensive, require high maintenance, and must comply with strict safety regulations.

:*Electrical heating units are unsafe compared to steam heating units. In steam systems, the physically steam controls the temperature, whereas in electrical heating units temperature is controlled by temperature controllers, which can fail or burn out.

The use of electricity carries with it some hazards depending on the environment. Extra care must be taken when using electrically-powered equipment in areas which may have combustible fluids, vapors, or dust, and where liquids may be present. (Seider 2009 pg 606)

==Conventional Power Station==

In general, most electricity is generated from a conventional coal-fired process, whether it be on-site or purchased from a provider. Coal-fired processes have been used to create electricity throughout history, and technological advances have increased its efficiency and use worldwide. In a coal-fired steam station—much like a nuclear station—water is turned into steam, which in turn drives turbine generators to produce electricity. There are several variations on how to create energy from coal, but here are the basics of how a coal-fired process works:

[[File:coalfired.png|thumb|border|center|middle|upright=4|link=|atl=|General Coal-Fired Process Diagram (Duke Energy Company)]]

*Heat is created:
::Before the coal is burned, it is pulverized to the fineness of talcum powder. It is then mixed with hot air and blown into the firebox of the boiler. Burning in suspension, the coal/air mixture provides the most complete combustion and maximum heat possible.

*Water turns to steam:
::Highly purified water, pumped through pipes inside the boiler, is turned into steam by the heat. The steam reaches temperatures of up to 1,000 degrees Fahrenheit and pressures up to 3,500 pounds per square inch, and is piped to the turbine.

*Steam turns the turbine:
::The enormous pressure of the steam pushing against a series of giant turbine blades turns the turbine shaft. The turbine shaft is connected to the shaft of the generator, where magnets spin within wire coils to produce electricity.

*Steam is converted back to water:
::After doing its work in the turbine, the steam is drawn into a condenser, a large chamber in the basement of the power plant. In this important step, millions of gallons of cool water from a nearby source (such as a river or lake) are pumped through a network of tubes running through the condenser. The cool water in the tubes converts the steam back into water that can be used over and over again in the plant.

*Repeat:
::The cooling water is returned to its source without any contamination, and the steam water is returned to the boiler to repeat the cycle.

===Advantages & Disadvantages of Coal-Fired Energy Production===

[[File:coalfiredtable.png|thumb|border|center|middle|upright=4|link=|atl=|Advantages and Disadvantages of Coal-Fired Processes (Seider 2008)]]

==Gas-Turbine Cogeneration Process==

When generating energy on-site, many plants use a gas-turbine cogeneration process. The thermal efficiency of a gas-turbine process is in the range of 70-80% while conventional power stations, such as coal-fired processes, have a 30-40% efficiency. The lower efficiency in more conventional power stations is attributed to wasted heat in the exhaust steam in the condenser. One example of a gas-turbine process is outlined in the following figure. Figure 3.1 is a gas-turbine cogeneration process with a heat recovery steam generator (waste-heat) boiler.

[[File:cogeneration.png|thumb|border|center|middle|upright=4|link=|atl=|Gas-Turbine Cogeneration with a heat recovery steam generator boiler (Towler 2012 Fig 3.1)]]

Overall, the process illustrated is not much different from a coal-fired process. The main differences are that the cogeneration process creates both electricity and a heat utility, and cogeneration processes use natural gas instead of coal. Many of the advantages and disadvantages are similar to those of the coal-fired process, but the cogeneration has a much higher efficiency, creates heat to be used in another process, and uses a more volatile and expensive fuel. The main advantage of cogeneration over coal-fired energy production is that heat is not wasted. In coal-fired processes, heat is released and wasted during electricity generation. Cogeneration captures some, if not all of the byproduct for heat, which is an extremely useful utility that will be discussed in subsequent sections. In summary, the cogeneration plant is superior to the coal-fired process because of its higher efficiency and ability to create a useful heat utility.

Obviously any engineer would design the cogeneration plant to meet at least the energy requirement necessary for plant operation, but cogeneration plants often times are designed to exceed the plant electricity requirement to drive another source of capital. Many describe this scenario as a "make or buy" scenario (Towler 2012). This scenario provides chemical producers leverage when negotiating contracts with outsourced electricity providers and this allows plants to purchase electricity at a wholesale price. This is a huge advantage of considering on-site electricity production because electricity is needed in relatively high quantities for all chemical plants. Being able to minimize electricity costs, or even profit off of electricity production is a huge economical consideration that all plants employ.

=Process Heating=
The key objective of this section is to discuss how processes are heated. Heating utilities are necessary for proper usage of distillers, reactors, condensers, and several other integral types of equipment. More specifically, steam, fired heat, and hot oil/specialized heat transfer fluids will be discussed in the following subsections.

==Steam==
Steam is the most commonly used heat utility used in chemical plants, and as a result understanding how it is used is essential in the study of Utility systems. Steam is used both as a process fluid (feedstock, diluent to absorb heat of reaction, heating agent, and stripping agent in absorbers and adsorbers ) and utility. It can be used to drive pumps and compressors, ejectors (for producing a vacuum), and heat exchangers. As one can clearly see, steam is a versatile, and useful utility.

Here are a few advantages of using steam as opposed to other methods of process heating (Towler 2012):

:*By controlling the pressure of the steam, one can control the temperature at which the heat is released. Having a strong control over the temperature is essential in several processes. =
:*Steam is an efficient heat source because the heat of condensation of steam is very high. Meaning that there is is high output per mass of utility at a constant temperature.
:*Heat exchangers that use steam are relatively cheap because condensing steam has a high heat transfer coefficient.
:*Steam is nonflammable, nontoxic, and inert to several process fluids.

Chemical plants generally have a network of pipelines exclusively for providing steam. These networks generally have steam at a low pressure, a medium pressure, and a high pressure. The image below illustrates a basic steam system.

[[File:Steam.png|thumb|border|center|middle|upright=4|link=|atl=|(Towler 2012 Fig 3.2)]]

In the diagram above, boiler feed water at a high pressure is preheated and fed to other boilers. These other boilers superheat the steam to create a high pressure and high temperature steam stream. The steam is superheated past the dew point to account for heat loss in the pipelines. A portion of the high pressure steam is used for process heating in areas of the plant that require high temperatures. The rest of the high pressure steam is turned into medium pressure steam by valves and steam turbines. The medium pressure steam is then used to heat medium temperature processes and to form low pressure steam. The low pressure steam can be used to heat low pressure processes and it can be expanded in condensing turbines to create shaft work and energy. In summary, steam can be used for an innumerable amount of action items in a plant. High pressure, medium pressure, and low pressure steam can all be used as a heat source. Low pressure steam has utilities in creating electricity and it also has several other uses.

==Fired Heat==

In many cases, processes in a plant require a heat source stronger than high pressure and temperature steam. That is when fired heat is used, which is generally at temperatures above 523K. Streams can be heated directly in the furnace tubes or via a hot oil circuit or heat transfer fluid, which will be discussed in detail in the next section. Most fired heaters use natural gas as fuel because it burns cleaner than fuel oil. A cleaner burning fuel is always advantageous due to environmental and safety concerns. Furthermore, natural gases usually result in less wear and tear in burners and fuel lines.

Depending on the application of the fired-heater, different design specifications can be implemented to make the fired-heater as efficient as possible. The basic construction of a fired heater starts with a cylindrical chamber that is lined with with refractory bricks. The stream that is to be heated flows through tubes inside of the furnace. These tubes can be arranged in several different arrangements such as, around the walls of the furnace, or in horizontal or vertical banks (Towler 2012). The figure below illustrates the basic construction of the fired-heater and varying tube arrangements.

[[File:firedheater.png|thumb|border|center|middle|upright=4|link=|atl=|Basic Fired-Heater Diagrams (Towler 2012 Figure 19.66)]]

Fuel is burned to heat the entire furnace, and the heat transfer occurs from the combustion gases inside of the furnace across the tubes that are filled with our desired stream. The heat transfer between the tubes and the furnace is accomplished primarily via radiation. Modern designs take advantage of convective heat transfer by adding a smaller section on top of the fired-heater where the combustion gases flow over banks of tubes as seen in (c) in the figure above. Heat transfer can be further improved via convection by adding fins or pins in the combustion section.

The cost of fired heat can be calculated by the cost of fuel fired. Natural gas and heating oil are traded as commodities and prices can be found at many online trading sites or business news sites (i.e., www.cnn.money.com). Past historic prices for forecasting can be found in the Oil and Gas Journal or from the U.S. Energy Information Adminstration (www.eia.gov).

===Fuels===
Fuel is burned in utility facilities such as boilers, electricity generation, and cogeneration, and can be in solid, liquid, or gas form. It can also be burned to provide heating for a process or stream or to drive pumps and compressors. The fuel is usually burned with excess air to ensure complete combustion.

A way of quantifying the amount of heat generated is by using the heating values. Higher heating value (HHV) and the lower heating value (LHV) are used. The heating is the total heat evolved by complete combustion of a fuel with dry air with both at 60 ⁰F and the flue gas after combustion brought back down to ⁰F. If the water vapor in the flue gas is not condensed, we obtain the LHV. If the water vapor is condensed, the value of heat evolved is a bit higher, and this is the HHV.
Heating values for solids and liquids are usually on a per-mass basis, and gases on a per-volume basis.
A specified amount of heating cannot be met with the amount of fuel calculated using only the HHV. There will be heat losses, the flue gas temperature will be greater than 60 ⁰F, and water in the flue gas will typically be vapor. (Seider 608)

==Hot Oil/Specialized Heat Transfer Fluids==

Specialized heat transfer fluids and hot oil circuits are used as heat sources when steam and fired heat is not appropriate. Specialized heat transfer fluids and hot oil circuits are extremely versatile in that they can be used in the temperature range of 323K to 673K. This range however is quite variable. For hot oils, the upper temperature limit is gauged based off of the thermal decomposition of the oil and coking/fouling of heat exchanger tubes.

Hot oil circuit systems are most commonly used when the plant has several small temperature heating requirements because it is more economically sound. Rather than having several fired heaters heat each small temperature requirement, it is much more economical to have one fired heater heat the hot oil and circulate that oil through each of the process to meet all of the heat needs. Hot oil systems are also generally favored over high pressure steam in processes that involve high pressure differentials between the process stream and high pressure steam. Hot oil systems are favored in this scenario because of safety concerns. If the steam were to leak, the pressure drop could cause serious safety issues.

Mineral oils are the most commonly used heat transfer fluids, and one prominent example is Dowtherm A. Dowtherm A is a combination of 26.5 wt% diphenyl in diphenyl oxide (Towler 2012) and is extremely thermally stable. These mineral oil systems generally require high flow rates.

=Process Cooling=

==Cooling Water==
Cooling water is used to cool and/or condense streams. Cooling water is usually circulated between process heat exchangers and a cooling tower. Water is cooled during downward motion by contact with air blown upwards, which can bring the water temperature to come within ~ 5 ⁰F of air’s wet-bulb temperature.Approximately 80% of the temperature reduction is due to evaporation of the cooling water and heat transfer to the surrounding air. Water can also be cooled in spray ponds and cooling ponds. Both work by providing high area for water to exchange heat with air.
Water in cooling towers is lost through drift and blowdown, and makeup is usually 1.5 to 3% of the circulating rate.
If a large natural body of water is nearby, it can be used as a source of cooling water and discharged downstream. This water is usually filtered to remove salts and impurities that may lead to fouling, but it is not treated.

==Refrigeration==
Cooling water can usually be used to cool a stream to ~ 100 ⁰F. Air can only cool to ~ 120 ⁰F. Air may be used in places where water is scarce or more costly to transport. To cool or condense streams below 100 ⁰F, a refrigerant is typically used. Chilled brine can also be used, but is less common.

Until 1995, CFC Freon R-12 (dichlorodifuloromethane) and HCFC Freon R-22 (chlorodifuloromethane) were commonly used refrigerants. However, the chlorine atom in the molecules caused the depletion of the ozone layer. The U.S. Clean Air Act Amendments of 1990 went into effect in 1995, and the production of these refrigerants has since stopped or been greatly reduced.

Cost estimates are based on ton-day of refrigeration, where a ton is the heat that needs to be removed to freeze 1 ton per day of water at 32 ⁰F. Substitutes have since been developed. R-134a is often used in place of R-12. According to the EPA, R-134a is not combustible at ambient conditions, and is essentially non-toxic under 400 ppm, and is not ozone-depleting. (Seider pg 607)

=Energy Efficiency=
One of the chief concerns in selecting and designing process utility systems for heating and cooling is how to achieve the most energy efficient design. There are countless means by which plants lose energy, two of the foremost being through the mixing of different temperature or pressure streams and through the disposal of warmed cooling water. (Seider, 2009) Proper utilities design can help mitigate each of these losses as well as many others. The energy efficiency of a plant will depend primarily on the heating and cooling methods that are being used and the overall system design itself. These two parameters are important in determining how well energy is transferred to the desired media as well as how well that energy is recovered and recycled.

==Hot Utility Efficiency==
As mentioned above, the most commonly used utilities for process heating in large scale processes are steam, fired heat, and hot oil heaters. Of these, steam is the most commonly used. Electricity, while efficient at creating power, is not a viable source of heat in large industrial processes. Common ranges of heating efficiency for these three methods are displayed in Table 1. (Towler and Sinnott, 2012; Broughton, 1994)

{| class="wikitable"
|+Table 1: Process Heating Efficiencies
|-
! Process Heating Method
! Typical Efficiency
|-
| Steam (via coal boiler)
| 72%
|-
| Steam (via gas boiler)
| 73%
|-
| Steam (via oil boiler)
| 75%
|-
| Fired Heat w/ Convective Section
| 85%
|-
| Fired Heat w/o Convective Section
| 60%
|-
| Hot Oil Heaters/Vaporizers
| 80-85%
|}

===Steam===
As steam is so popular for heating purposes, it is useful to analyze the numerous ways in which losses can occur in steam systems. There are five primary sources of inefficiency and heat loss in the generation and distribution of steam throughout a process plant:

*The heat content of boiler exhaust gas
*Incomplete combustion of boiler fuel
*Radiant losses from the boiler exterior
*Blowdown losses
*Distribution losses (pipe transport, steam traps, etc.)

The first four of these losses take place at the boiler and contribute to the heating efficiencies seen in Table 1 for steam created with coal, gas, and oil. Several methods can be used to minimize these losses, one of the foremost being the control of air-to-fuel ratio in the boiler. This ratio must be managed by weighing losses due to uncombusted fuel against losses due to the heat content of excess exhaust gas. (Broughton, 1994)

[[File:Air-fuel ratio.PNG|thumb|border|center|middle|upright=4|link=|atl=|Air-to-fuel ratio relationship with heat losses (Broughton Fig 2.3)]]

The air-to-fuel ratio can be optimized using a feedback process controller. The control system will analyze the oxygen content of exhaust air and adjust the incoming air flow rate to achieve a set percentage of excess air. While desired excess oxygen will vary depending on the type of fuel, it is consistently seen that in the minimum loss-range a 1% increase in excess air will result in a 1% decrease in efficiency. ("Oxygen Control")

In addition to air-to-fuel ratio management, steam heat losses at the boiler can be mitigated via energy recovery, which is discussed in further detail [[Utility_systems#Energy_Recovery|below]].

Outside the boiler, losses in distribution of steam throughout facilities can decrease energy efficiency by a significant amount. This can cause up to a 60% increase in losses, but typically results in an overall steam efficiency of 50-55% (down from approximately 75% at the boiler). (Broughton, 1994) There are several ways that this issue can be addressed. First and foremost are steam metering systems, which can be used to monitor heat supply to separate sections of the process facility, check efficiency of fuel use, and determine which processes in a given plant have the highest steam requirements. Another method of minimizing transport losses is to decentralize steam generation systems. It can be advantageous to have numerous smaller boilers rather than a single centralized boiler so that steam does not have to travel as long of distances to reach its destination.

==Cold Utility Efficiency==
Efficiency in cooling processes depends more on the method used, and by extension the amount of coolant needed. Water and air utility efficiencies depend primarily on the fluid flow required to maintain the system at a desired temperature, while powered refrigeration utilities (for colder processes) have efficiencies at approximately 60%--but ranging up to 90%--of Carnot cycle efficiency, a metric of ideal refrigeration efficiency. (Towler and Sinnott, 2012) Cooling systems represent by definition a loss of energy from the main process to the utility stream, and as such it is often useful to find other uses for the heated media before discharge.

==Energy Recovery==
Recovery and recycle of energy is perhaps the most important aspect of creating an energy efficient plant design, and it is important for process engineers to fully consider possibilities for heat recovery in order to aid in economic viability.

===Process Heat Exchange===
Heat exchanger networks are a very common energy recovery method in industrial processes. These networks most frequently allow energy from heated product streams to be transferred to feed streams that must be brought up to process temperature. (Biegler, 1997) More information on the function and design of heat exchanger networks can be found on the [[Heat_exchanger#Heat_Exchanger_Networks|heat exchanger wiki page]]. The following are several examples of energy recovery via heat exchange that are used in industrial processes.

In distillation columns the bottoms and distillate effluents have the potential for energy exchange. Though the condenser at the top of the column cannot supply its waste heat to the reboiler due to their respective temperatures, the effluent streams can supply heat to the feed via a feed-effluent exchanger. This reduces the utility requirements to raise the feed to column temperature. (Biegler, 1997)

[[File:Feed_sterilization.jpg|thumb|border|right|middle|upright=4|link=|atl=|Feed sterilization schematic. (Towler and Sinnott Fig. 3.30)]]

Feed sterilization, commonly used in the food industry, is a common application for heat recovery through process stream heat exchange. In this application, the feed must be heated for a certain amount of time to kill any biological contaminants, after which it can be used to heat the new raw feed for sterilization. This reduces energy demands on the steam heater and thus reduces cost. (Towler and Sinnott, 2012)

In multi-vessel batch processes it can be advantageous to exchange heat as the process fluid is being transferred between vessels. Like the previous examples, this reduces the utility needed to bring the colder feed up to process temperature, thus reducing costs. (Towler and Sinnott, 2012)

[[File:Batch_heat_exchange.jpg|thumb|border|center|middle|upright=4|link=|atl=|Multi-vessel batch heat exchange schematic (Towler and Sinnott Fig. 3.31)]]

===Utility Regeneration===
[[File:Waste_heat_boiler.jpg|thumb|border|right|middle|200px|upright=4|link=|atl=|Industrial modular HRSG]]
When recovery of waste heat via transfer to other process streams is inconvenient or impossible, energy efficiency can still be improved through the regeneration of utilities. This is commonly done through the regeneration of steam by removing heat from exiting streams or from highly exothermic reactions. Waste heat in exiting streams can be removed via heat recovery steam generators (HRSGs), and is most often used on exiting gas streams. Heat recovery from reactions is a viable option when the reactor temperature will be at 150 C or above, as this will create steam at high enough pressure to be used in other processes. (Towler and Sinnott, 2012)

In the case of steam, waste heat and water treatment losses can be recovered from the utility generation process itself. One of the most common ways to do this is with an economizer. As seen in the schematic below, economizers have heat exchangers fit to the exhaust gas flow in order to transfer waste heat from these gases to the incoming boiler feed water. This can result in fuel energy savings of approximately 15% for typical excess air percentages in steam boilers. (Broughton, 1994; "Economizers")

[[File:Economizer.png|thumb|border|center|middle|upright=4|link=|atl=|Economizer schematic ("Economizers") and economizer energy recovery correlations (Broughton Fig 2.2)]]

[[File:Condensate recycle.PNG|thumb|border|right|middle|200px|upright=4|link=|atl=|Condensate return fuel savings (Broughton Fig 2.4)]]

Additionally, there are still more opportunities for recovery in steam generation processes. One typical case is the use of waste heat to preheat air entering the boiler. Optimistically, this can result in a 5% improvement in heat recovery for the system. Another method is the recycle of steam condensate back to the boiler. This has two main benefits: it retains the heat still present in the condensed steam and it retains the pretreatment chemicals added to the steam. (Broughton, 1994)

Furthermore, there is opportunity for energy recovery in the expansion of compressed gas through a turbine to create electricity, a process that can be economically viable given sufficiently high flows or pressure. One of the primary application of this type of energy recovery is in the creation of medium pressure and low pressure steam. In most processes, all steam is generated as high pressure steam and can be expanded through a turbine to decrease its pressure. Such technology has also been used in processes to synthesize ammonia, perform air separations, and synthesize nitric acid. (Towler and Sinnott, 2012) Recently, however, there has been a particularly strong interest for energy recovery in the natural gas industry, when gas is decompressed from major pipelines to residential low-pressure pipelines. A 2001 study estimated that there is the potential to recover 21 TWh, representing 11% of natural gas transport energy, via gas expansion. (Lehman)

=Process water and boiler-feed water=
Process water is water that will be directly used in the process. Boiler-feed water (BFW) is used to produce steam. Both may need to be purified to prevent impurities from contaminating a process or from foul equipment. It can be used as a cooling stream when the temperature of the stream to be cooled is greater than ~300 ⁰F. Cost of BFW can be partially offset by the steam credit.

Process water that undergoes moderate pretreatment can cost ~ $0.75/1,000 gal.

Extensive treatment ~ $6.00/1,000 gal.

Sterilized for pharmaceutical processes ~ $550/1,000 gal. (Seider pg 608)

=Demineralized Water=
In demineralized water, minerals have been removed by ion exchange. In boiler feed water, this is to prevent salt deposition, corrosion, formation of foam, and sluicing. In process water, the ions may contaminate the process.

=Waste Treatment=
Most chemical processes will produce some sort of waste. Disposal occurs to the atmosphere (in the case of some gases), sewers, body of water, or a landfill. Waste may require some treatment before disposal to meet regulations. Depending on process economics, byproducts may be recovered and processed. (Seider 2009 pg 609)

==Wastewater Treatment==
The United States EPA regulates industrial wastewater disposal through the Clean Water Act, introduced in 1948 as the Federal Water Pollution Control Act and amended to its current form in 1972. The sweeping 1972 amendments allowed the EPA to prevent industries and persons from discharging contaminated water into fresh water sources and set water quality standards. (Summary of the Clean Water Act) In accordance with this law, process plants in the United States treat wastewater at on-site or near-site treatment centers before releasing it into the surrounding environment.

Wastewater effluent streams, along with water runoff from around the plant, are treated to control for pH, toxicity, suspended solids, and biological oxygen demand (for aquatic life protection) prior to discharge. Each of these controls is typically addressed with a separate method. Acidity and basicity is balanced through the addition of an acid or alkaline solution. Toxic wastewater may be treated with chemical processes or simply diluted to safe concentrations. Suspended solids can be removed via filtration and/or with clarifiers. Oxygen demand of wastewater can be mitigated using activated sludge treatment processes. Once the water quality complies with the EPA, and state-mandated, regulations, it can be safely released. More information on the large number of industry-specific guidelines for waste effluent can be found on the EPA website (http://www.epa.gov/eg/industrial-effluent-guidelines).

=Air-Pollution Management=
===Introduction===

In the United States air pollution is regulated in the Clean Air Act, and almost all pollutant emitting plants are regulated under this law. The types of plants that can release significant emissions include petroleum refineries, sulfur recovery plants, carbon-black plants, fuel conversion plants, chemical process plants, fossil fuel plants, and petroleum storage and transfer facilities. To receive permission to construct a plant must undergo a review to show that it will not cause a violation of the Ambient Air Quality Standards(Peters, 2003).

==Methods==

There are two major types of pollutants that are released into the air, particulates and and gaseous pollutants. Particulates can be removed with mechanical forces while gaseous pollutants typically need to removed by chemical or physical means (Peters, 2003).

Particulates and volatile pollutants that need to be removed before disposal may be present. Particle removal equipment includes: cyclones, wet scrubbers, electrostatic precipitators, and fabric-filter systems (Seider 2009) such as bag filters (. The two charts below are from ''Plant Design and Economics for Chemical Engineers'' and show the types of equipment, separation methods, and particle sizes in different pollutant separation technologies (Peters, 1991).

[[File:AirPolChart.png|thumb|border|center|middle|upright=4|link=|atl=|Different Methods of Particle Separation, the Particle Sizes they Can Remove, and the Technologies Used]]

[[File:AirPolTable.png|thumb|border|center|middle|upright=4|link=|atl=| This Chart Shows the Specifications and Limitations of Different Separations Technologies Including Particle Size, Efficiencies, and Temperatures(Peters, 1991)]]

===Wet Scrubbers===
Wet scrubber use lime or limestone and water to remove SO2 and acid gases. The mixture can be injected into a scrubber or the gases can be bubbled through this mixture. This results in removal of 90-98% of SO2 and and acid gases (Clean Coal Technologies).

===Dry Scrubbers===
Dry scrubber blow powdered adsorbents into a vessel with gases and then after it has captured the SO2 and acid gases it is separated from gas using a fabric filter. These systems remove 90-93% of the contaminants (Clean Coal Technologies).

===Low-NOx Burners===
The purpose of low NOx burners is to decrease the amount of NOx created when the coal is burned. This is done by injecting coal and air in boilers. This can result in 40-50% NOx reduction. If air is injected into the area above the burner this can actually cause almost 70% NOx reduction (Clean Coal Technologies).

===Selective Non-Catalytic and Catalytic Reductions===
These systems inject ammonia into gases to remove NOx. The Catalytic reductions add a catalyst to the ammonia being injected to aid in the removal. Con-catalytic reductions result in about 35% removal, but adding a catalyst can increase that amount to about 90%. The catalytic-reduction can remove up to 80% of mercury as well (Clean Coal Technologies).

===Fabric Filters===
Also known as baghouses,these [[Separation processes#Filtration| filters]] remove particulates by passing air through filters. These can separate as much as 99.9% of particulate matter (Clean Coal Technologies).

===Electrostatic Precipitators===
Electrostatic Precicpitators remove particulate matter as gas passes through a device that has charged metal plates. The particles are then removed because of static electricity. These systems can remove between 99-99.9% of particulate matter (Clean Coal Technologies).

===Super-critical Boilers===
Super-Critical Boilers and Ultra-Supercritical Boilers operate at temperature and pressure higher than regular boilers. By operating at higher temperature these systems become more efficient. Super Critical Boilers typically have 10%-20% CO2 emissions than other similar sub-critical technologies. Ultra-Supercritical boilers can be as much as 30% more efficent than sub-critical technologies (Clean Coal Technologies).

===More Information===

For more information regarding some of these separations equipment see [[Solids-involved equipment]]

For more information Regarding Cyclones see [[Separation processes#Cyclones| Cyclones]], and for modeling cyclones see their[[Solids-involved equipment#HYSYS Simulation|HYSYS Simulation]]

Methods for removing inorganic and organic gaseous pollutants include: absorption, adsorption, condensation, and combustion (Seider 2009 pg 609). A list of typical of gases pollutants and their sources from ''Plant Design and Economics for Chemical Engineers'' is shown below (Peters, 1991).

[[File:GaseousPollutants.png|thumb|border|center|middle|upright=4|link=|atl=| Common Gaseous Pollutants and their Sources]]

Typically Gas-liquid absorption processes are done completed in a vertical, countercurrent, flow through packed, plate, or spray towers. These systems require good liquid-gas contact and proper equipment. These systems also often have significant energy consumption because of large pressure drops (Peters, 2003). For high volume systems absorption by scrubbing with water or another solvent is the most widely used method (Towler, 2012). Dry adsorbents can be used to remove the last races of gaseous pollutants. Adsorption typically requires blowers, condensers, separators, and controls. You also typically need two packed beds so that one can be used while the other is regenerated. Examples of adsorbents are molecular sieves and activated carbon. Incineration is typically used when there are gas streams that have no recovery value. This can be done with direct flame or catalytic oxidation. Catalytic oxidation usually has higher capital costs, but lower operating costs because it does not require fuel.

You can find more information on [[Separation processes#Absorption| Absorption]] and [[Separation processes#Adsorption| Adsorption]] in [[Separation processes]]

===Outcomes===

====United States Common Pollutant Emissions====

The United States implemented the Clean Air Act in 1970 and since then emissions in the U.S. have been drastically reduced. As a matter of fact despite increases in the population in the last 35 years the amount of pollutants emitted have decreased by almost 70% (EPA). More importantly this demonstrates that reasonable efforts can be put towards environmental protection without causing too much harm to industry.

[[File:EPA.png|thumb|border|center|middle|upright=4|link=|atl=| GDP and Other Growth Factors Vs Common Pollutants Released]]

===China Clean Coal===

====Success Through 2005====

Coal is a very inexpensive and abundant source of energy and is abundant in China (Xu 2010). In China Coal the cause of 90% of SO2 emissions, 70% of dust emissions, and 67% of NOx emissions, and 70% of CO2 emissions. These numbers are staggering considering that Coal has never been more than 50 percent of China's Energy Supply (Xu 2010).

[[File:Coal.jpg|thumb|border|center|middle|upright=4|link=|atl=| Coal Consumption for Power and the Percentage of Coal Used (Xu 2010)]]

Despite the increasing coal consumption, high efficiency electric dust removal systems with efficiencies that are as high as 99.6% have greatly decreased soot emissions to 32% below 1980's levels as of 2005.

[[File:Soot.jpg|thumb|border|center|middle|upright=4|link=|atl=| SO2 and Soot Emissions from 1981 to 2005 (Xu 2010)]]

====Technologies and Life Cycle Comparison====
As of 2010 China was consuming 48.2% of coal globally. Four potential technologies integrated gasification combined cycle (IGCC), sub-critical coal power generation (Sub-C); super-critical coal power generation (Super-C) ultrasuper-critical coal power generation (USC). These technologies are compared mostly on the basis of net generating efficiency and efficiency. Net generating efficiency is the output of the plant divided by the total available energy in the fuel used (Liang, 2013).

where efficiency is define as: η=E/Etotal

and

Etotal=Emining+Etransportation+Egeneration

[[File:technologies.PNG|thumb|border|center|middle|upright=4|link=|atl=|A comparison of many of the parameters important to lfe-cycle analysis (Liang, 2013)]]

This demonstrates that USC has the best net generation efficiency, while also having the largest capacity for a single system.

[[File:Lifecycle2.PNG|thumb|border|center|middle|upright=4|link=|atl=| (Liang, 2013)]]

This breaks down the energy efficiency of each system and demonstrates that USC is the most efficient

[[File:Lifecycle.PNG|thumb|border|center|middle|upright=4|link=|atl=| (Liang, 2013)]]

This demonstrates that the capital cost per unit of energy production is very high for IGCC, but shows that Super-C and USC are very competitive.

[[File:Price.PNG|thumb|border|center|middle|upright=4|link=|atl=| Capital cost in $/kW of clean coal power-generation technologies (Liang, 2013)]]

Lastly IGCC and USC are lowest on different element, but since these all have different global warming potential it is difficult to tell which is the most efficient.

And the table below defines all the relevant terms.

[[File:GWP.PNG|thumb|border|center|middle|upright=4|link=|atl=| Environemntal measures and their explanations (Liang, 2013)]]

All of the data suggests that USC is the highest in both energy efficiency and net generating efficiency while competitive in price and the lowest on emissions among the different boilers.

=Solid Waste=
U.S. federal regulations require that solid waste be classified as hazardous or nonhazardous. Conditions for a classification of hazardous include: ignitability, corrosivity, reactivity, toxicity, or posing a substantial threat to the surrounding environment and its inhabitants. Hazardous waste must be treated on- or near-site before being removed in containers. Non-hazardous waste may be landfilled or incinerated in some cases. A typical estimate of costs for waste disposal is $0.03/lb for nonhazardous solids and $0.10/lb for hazardous solids. (Seider 2009 pg 609)

=References=

*Biegler LT, Grossmann IE, Westerberg AW. ''Systematic Methods of Chemical Process Design'' Prentice-Hall: Upper Saddle River, 1997.

*Broughton, J. ''Process Utility Systems: Introduction to Design, Operation and Maintenance'' Institution of Chemical Engineers: Rugby, Warwickshire, UK, 1994.

*Clean Coal Technologies. America's Coalition for Clean Coal Electricity. http://www.americaspower.org/clean-coal-technologies-1663/

*Duke Energy Company (2013). How Do Coal Fired Plants Work? Charlotte: Duke Energy.

*"Economizers." Energy Solutions Center Inc.: Boiler Consortium. http://www.cleanboiler.org/Eff_Improve/Efficiency/Economizers.asp. Accessed February 21, 2016.

*Lehman B, Worrell E. ''Electricity Production from Natural Gas Pressure Recovery Using Expansion Turbines.'' Lawrence Berkeley National Laboratory; 2001.

*"Overview of the Clean Air Act and Air Pollution." Environmental Protection Agency. November 17, 2015. http://www.epa.gov/clean-air-act-overview. Accessed February 5, 2016.

*"Oxygen Control." Energy Solutions Center Inc.: Boiler Consortium. http://www.cleanboiler.org/Eff_Improve/Efficiency/Oxygen_Control.asp. Accessed February 21, 2016.

*Peters, Max S.; Timmerhaus, Klaus D.; West, Ronald E. (2003). " Plant Design and Economics for Chemical Engineers." McGraw Hill Higher Education.

*Ruina Xu Energy (2010). "Clean coal technology development in China." Elsevier.

*Seider, Seader, Lewin, Widagdo. (2009). ''Plant Design and Economics for Chemical Engineers, 5th Edition.'' Hoboken: Wiley.

*Seider, Seader, Lewin. (2008). ''Product and Process Design Principles, 2nd Edition.'' Hoboken: Wiley.

*Summary of the Clean Water Act. United States EPA website. http://www.epa.gov/laws-regulations/summary-clean-water-act

*Towler, G.P. and Sinnot, R. (2012). ''Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design.'' Elsevier.

*Turton R.; Bailie, R.C.; Whiting, W.B.; Shaeiwitz J.A.; Bhattacharyya D. (2012). ''Analysis, Synthesis, and Design of Chemical Processes.'' Upper Saddle River: Prentice Hall.

*G.D. Ulrich, A Guide to Chemical Engineering Process Design and Economics, Wiley: New York, 1984.

*Xiaoye Liang, Zhihua Wang, Zhijun Zhou, , Zhenyu Huang, Junhu Zhou, and Kefa Cen. (2013). "Up-to-date life cycle assessment and comparison study of clean coal power generation technologies in China" Elsevier.

Utility systems

2016-02-22T05:01:24Z

Bts511: /* Wet Scrubbers */

Authors: David Chen, [2014] Joshua Lee, [2015] Brett Sleyster, [2016] and Tom Aunins [2016] 

Stewards: David Chen, Jian Gong, and Fengqi You

Date Presented: January 13, 2014 /Date Revised: January 14, 2014

 


__TOC__

=Introduction=
Many chemical processes do not take place at ambient temperature or pressures. In order to reach these non-ambient conditions, utilities will have to be used to raise or lower temperatures and compress gases. (Towler, Towler/UOP) Utilities often contribute 5 to 10% of the price of a product, and may come from public or private utility companies or on-site plants. For purchased utilities, costs depend on consumption, while for company-owned utilities, there will be both capital and operating costs. They include things such as steam for heating, electricity, cooling water, refrigeration, fuels such as natural gas, wastewater treatment, waste disposal, and landfill. Steam is often the largest utility cost. Cogeneration unit can supply electricity accompanied with different steam pressures. (Seider 2009)

=Electricity=

Electricity is used to power many different kinds of equipment. It has many advantages: it is efficient (> 90%), reliable, available in a wide range of power, shaft speeds, designs, lifetimes, convenience, costs, and maintenance. It is generally used up to 200 hp, and sometimes over 10,000 Hp. In chemical process plants, the electricity demand is generally determined by the work or energy required for compression, pumping, air cooling, lights, and many other items. This electricity often times is purchased from local electricity providers, but many plants generate their own electricity via sophisticated processes.

Electricity is rarely used as a primary heat utility in large-scale chemical plants for a variety of reasons. The main disadvantages of using electricity as a heat utility are as follows (Towler 2012):

:*Heat from electricity is two to three times more expensive than heat from fuels. This is attributed to the lack of efficiency when creating heat from electricity.

:*Electrical heating units are expensive, require high maintenance, and must comply with strict safety regulations.

:*Electrical heating units are unsafe compared to steam heating units. In steam systems, the physically steam controls the temperature, whereas in electrical heating units temperature is controlled by temperature controllers, which can fail or burn out.

The use of electricity carries with it some hazards depending on the environment. Extra care must be taken when using electrically-powered equipment in areas which may have combustible fluids, vapors, or dust, and where liquids may be present. (Seider 2009 pg 606)

==Conventional Power Station==

In general, most electricity is generated from a conventional coal-fired process, whether it be on-site or purchased from a provider. Coal-fired processes have been used to create electricity throughout history, and technological advances have increased its efficiency and use worldwide. In a coal-fired steam station—much like a nuclear station—water is turned into steam, which in turn drives turbine generators to produce electricity. There are several variations on how to create energy from coal, but here are the basics of how a coal-fired process works:

[[File:coalfired.png|thumb|border|center|middle|upright=4|link=|atl=|General Coal-Fired Process Diagram (Duke Energy Company)]]

*Heat is created:
::Before the coal is burned, it is pulverized to the fineness of talcum powder. It is then mixed with hot air and blown into the firebox of the boiler. Burning in suspension, the coal/air mixture provides the most complete combustion and maximum heat possible.

*Water turns to steam:
::Highly purified water, pumped through pipes inside the boiler, is turned into steam by the heat. The steam reaches temperatures of up to 1,000 degrees Fahrenheit and pressures up to 3,500 pounds per square inch, and is piped to the turbine.

*Steam turns the turbine:
::The enormous pressure of the steam pushing against a series of giant turbine blades turns the turbine shaft. The turbine shaft is connected to the shaft of the generator, where magnets spin within wire coils to produce electricity.

*Steam is converted back to water:
::After doing its work in the turbine, the steam is drawn into a condenser, a large chamber in the basement of the power plant. In this important step, millions of gallons of cool water from a nearby source (such as a river or lake) are pumped through a network of tubes running through the condenser. The cool water in the tubes converts the steam back into water that can be used over and over again in the plant.

*Repeat:
::The cooling water is returned to its source without any contamination, and the steam water is returned to the boiler to repeat the cycle.

===Advantages & Disadvantages of Coal-Fired Energy Production===

[[File:coalfiredtable.png|thumb|border|center|middle|upright=4|link=|atl=|Advantages and Disadvantages of Coal-Fired Processes (Seider 2008)]]

==Gas-Turbine Cogeneration Process==

When generating energy on-site, many plants use a gas-turbine cogeneration process. The thermal efficiency of a gas-turbine process is in the range of 70-80% while conventional power stations, such as coal-fired processes, have a 30-40% efficiency. The lower efficiency in more conventional power stations is attributed to wasted heat in the exhaust steam in the condenser. One example of a gas-turbine process is outlined in the following figure. Figure 3.1 is a gas-turbine cogeneration process with a heat recovery steam generator (waste-heat) boiler.

[[File:cogeneration.png|thumb|border|center|middle|upright=4|link=|atl=|Gas-Turbine Cogeneration with a heat recovery steam generator boiler (Towler 2012 Fig 3.1)]]

Overall, the process illustrated is not much different from a coal-fired process. The main differences are that the cogeneration process creates both electricity and a heat utility, and cogeneration processes use natural gas instead of coal. Many of the advantages and disadvantages are similar to those of the coal-fired process, but the cogeneration has a much higher efficiency, creates heat to be used in another process, and uses a more volatile and expensive fuel. The main advantage of cogeneration over coal-fired energy production is that heat is not wasted. In coal-fired processes, heat is released and wasted during electricity generation. Cogeneration captures some, if not all of the byproduct for heat, which is an extremely useful utility that will be discussed in subsequent sections. In summary, the cogeneration plant is superior to the coal-fired process because of its higher efficiency and ability to create a useful heat utility.

Obviously any engineer would design the cogeneration plant to meet at least the energy requirement necessary for plant operation, but cogeneration plants often times are designed to exceed the plant electricity requirement to drive another source of capital. Many describe this scenario as a "make or buy" scenario (Towler 2012). This scenario provides chemical producers leverage when negotiating contracts with outsourced electricity providers and this allows plants to purchase electricity at a wholesale price. This is a huge advantage of considering on-site electricity production because electricity is needed in relatively high quantities for all chemical plants. Being able to minimize electricity costs, or even profit off of electricity production is a huge economical consideration that all plants employ.

=Process Heating=
The key objective of this section is to discuss how processes are heated. Heating utilities are necessary for proper usage of distillers, reactors, condensers, and several other integral types of equipment. More specifically, steam, fired heat, and hot oil/specialized heat transfer fluids will be discussed in the following subsections.

==Steam==
Steam is the most commonly used heat utility used in chemical plants, and as a result understanding how it is used is essential in the study of Utility systems. Steam is used both as a process fluid (feedstock, diluent to absorb heat of reaction, heating agent, and stripping agent in absorbers and adsorbers ) and utility. It can be used to drive pumps and compressors, ejectors (for producing a vacuum), and heat exchangers. As one can clearly see, steam is a versatile, and useful utility.

Here are a few advantages of using steam as opposed to other methods of process heating (Towler 2012):

:*By controlling the pressure of the steam, one can control the temperature at which the heat is released. Having a strong control over the temperature is essential in several processes. =
:*Steam is an efficient heat source because the heat of condensation of steam is very high. Meaning that there is is high output per mass of utility at a constant temperature.
:*Heat exchangers that use steam are relatively cheap because condensing steam has a high heat transfer coefficient.
:*Steam is nonflammable, nontoxic, and inert to several process fluids.

Chemical plants generally have a network of pipelines exclusively for providing steam. These networks generally have steam at a low pressure, a medium pressure, and a high pressure. The image below illustrates a basic steam system.

[[File:Steam.png|thumb|border|center|middle|upright=4|link=|atl=|(Towler 2012 Fig 3.2)]]

In the diagram above, boiler feed water at a high pressure is preheated and fed to other boilers. These other boilers superheat the steam to create a high pressure and high temperature steam stream. The steam is superheated past the dew point to account for heat loss in the pipelines. A portion of the high pressure steam is used for process heating in areas of the plant that require high temperatures. The rest of the high pressure steam is turned into medium pressure steam by valves and steam turbines. The medium pressure steam is then used to heat medium temperature processes and to form low pressure steam. The low pressure steam can be used to heat low pressure processes and it can be expanded in condensing turbines to create shaft work and energy. In summary, steam can be used for an innumerable amount of action items in a plant. High pressure, medium pressure, and low pressure steam can all be used as a heat source. Low pressure steam has utilities in creating electricity and it also has several other uses.

==Fired Heat==

In many cases, processes in a plant require a heat source stronger than high pressure and temperature steam. That is when fired heat is used, which is generally at temperatures above 523K. Streams can be heated directly in the furnace tubes or via a hot oil circuit or heat transfer fluid, which will be discussed in detail in the next section. Most fired heaters use natural gas as fuel because it burns cleaner than fuel oil. A cleaner burning fuel is always advantageous due to environmental and safety concerns. Furthermore, natural gases usually result in less wear and tear in burners and fuel lines.

Depending on the application of the fired-heater, different design specifications can be implemented to make the fired-heater as efficient as possible. The basic construction of a fired heater starts with a cylindrical chamber that is lined with with refractory bricks. The stream that is to be heated flows through tubes inside of the furnace. These tubes can be arranged in several different arrangements such as, around the walls of the furnace, or in horizontal or vertical banks (Towler 2012). The figure below illustrates the basic construction of the fired-heater and varying tube arrangements.

[[File:firedheater.png|thumb|border|center|middle|upright=4|link=|atl=|Basic Fired-Heater Diagrams (Towler 2012 Figure 19.66)]]

Fuel is burned to heat the entire furnace, and the heat transfer occurs from the combustion gases inside of the furnace across the tubes that are filled with our desired stream. The heat transfer between the tubes and the furnace is accomplished primarily via radiation. Modern designs take advantage of convective heat transfer by adding a smaller section on top of the fired-heater where the combustion gases flow over banks of tubes as seen in (c) in the figure above. Heat transfer can be further improved via convection by adding fins or pins in the combustion section.

The cost of fired heat can be calculated by the cost of fuel fired. Natural gas and heating oil are traded as commodities and prices can be found at many online trading sites or business news sites (i.e., www.cnn.money.com). Past historic prices for forecasting can be found in the Oil and Gas Journal or from the U.S. Energy Information Adminstration (www.eia.gov).

===Fuels===
Fuel is burned in utility facilities such as boilers, electricity generation, and cogeneration, and can be in solid, liquid, or gas form. It can also be burned to provide heating for a process or stream or to drive pumps and compressors. The fuel is usually burned with excess air to ensure complete combustion.

A way of quantifying the amount of heat generated is by using the heating values. Higher heating value (HHV) and the lower heating value (LHV) are used. The heating is the total heat evolved by complete combustion of a fuel with dry air with both at 60 ⁰F and the flue gas after combustion brought back down to ⁰F. If the water vapor in the flue gas is not condensed, we obtain the LHV. If the water vapor is condensed, the value of heat evolved is a bit higher, and this is the HHV.
Heating values for solids and liquids are usually on a per-mass basis, and gases on a per-volume basis.
A specified amount of heating cannot be met with the amount of fuel calculated using only the HHV. There will be heat losses, the flue gas temperature will be greater than 60 ⁰F, and water in the flue gas will typically be vapor. (Seider 608)

==Hot Oil/Specialized Heat Transfer Fluids==

Specialized heat transfer fluids and hot oil circuits are used as heat sources when steam and fired heat is not appropriate. Specialized heat transfer fluids and hot oil circuits are extremely versatile in that they can be used in the temperature range of 323K to 673K. This range however is quite variable. For hot oils, the upper temperature limit is gauged based off of the thermal decomposition of the oil and coking/fouling of heat exchanger tubes.

Hot oil circuit systems are most commonly used when the plant has several small temperature heating requirements because it is more economically sound. Rather than having several fired heaters heat each small temperature requirement, it is much more economical to have one fired heater heat the hot oil and circulate that oil through each of the process to meet all of the heat needs. Hot oil systems are also generally favored over high pressure steam in processes that involve high pressure differentials between the process stream and high pressure steam. Hot oil systems are favored in this scenario because of safety concerns. If the steam were to leak, the pressure drop could cause serious safety issues.

Mineral oils are the most commonly used heat transfer fluids, and one prominent example is Dowtherm A. Dowtherm A is a combination of 26.5 wt% diphenyl in diphenyl oxide (Towler 2012) and is extremely thermally stable. These mineral oil systems generally require high flow rates.

=Process Cooling=

==Cooling Water==
Cooling water is used to cool and/or condense streams. Cooling water is usually circulated between process heat exchangers and a cooling tower. Water is cooled during downward motion by contact with air blown upwards, which can bring the water temperature to come within ~ 5 ⁰F of air’s wet-bulb temperature.Approximately 80% of the temperature reduction is due to evaporation of the cooling water and heat transfer to the surrounding air. Water can also be cooled in spray ponds and cooling ponds. Both work by providing high area for water to exchange heat with air.
Water in cooling towers is lost through drift and blowdown, and makeup is usually 1.5 to 3% of the circulating rate.
If a large natural body of water is nearby, it can be used as a source of cooling water and discharged downstream. This water is usually filtered to remove salts and impurities that may lead to fouling, but it is not treated.

==Refrigeration==
Cooling water can usually be used to cool a stream to ~ 100 ⁰F. Air can only cool to ~ 120 ⁰F. Air may be used in places where water is scarce or more costly to transport. To cool or condense streams below 100 ⁰F, a refrigerant is typically used. Chilled brine can also be used, but is less common.

Until 1995, CFC Freon R-12 (dichlorodifuloromethane) and HCFC Freon R-22 (chlorodifuloromethane) were commonly used refrigerants. However, the chlorine atom in the molecules caused the depletion of the ozone layer. The U.S. Clean Air Act Amendments of 1990 went into effect in 1995, and the production of these refrigerants has since stopped or been greatly reduced.

Cost estimates are based on ton-day of refrigeration, where a ton is the heat that needs to be removed to freeze 1 ton per day of water at 32 ⁰F. Substitutes have since been developed. R-134a is often used in place of R-12. According to the EPA, R-134a is not combustible at ambient conditions, and is essentially non-toxic under 400 ppm, and is not ozone-depleting. (Seider pg 607)

=Energy Efficiency=
One of the chief concerns in selecting and designing process utility systems for heating and cooling is how to achieve the most energy efficient design. There are countless means by which plants lose energy, two of the foremost being through the mixing of different temperature or pressure streams and through the disposal of warmed cooling water. (Seider, 2009) Proper utilities design can help mitigate each of these losses as well as many others. The energy efficiency of a plant will depend primarily on the heating and cooling methods that are being used and the overall system design itself. These two parameters are important in determining how well energy is transferred to the desired media as well as how well that energy is recovered and recycled.

==Hot Utility Efficiency==
As mentioned above, the most commonly used utilities for process heating in large scale processes are steam, fired heat, and hot oil heaters. Of these, steam is the most commonly used. Electricity, while efficient at creating power, is not a viable source of heat in large industrial processes. Common ranges of heating efficiency for these three methods are displayed in Table 1. (Towler and Sinnott, 2012; Broughton, 1994)

{| class="wikitable"
|+Table 1: Process Heating Efficiencies
|-
! Process Heating Method
! Typical Efficiency
|-
| Steam (via coal boiler)
| 72%
|-
| Steam (via gas boiler)
| 73%
|-
| Steam (via oil boiler)
| 75%
|-
| Fired Heat w/ Convective Section
| 85%
|-
| Fired Heat w/o Convective Section
| 60%
|-
| Hot Oil Heaters/Vaporizers
| 80-85%
|}

===Steam===
As steam is so popular for heating purposes, it is useful to analyze the numerous ways in which losses can occur in steam systems. There are five primary sources of inefficiency and heat loss in the generation and distribution of steam throughout a process plant:

*The heat content of boiler exhaust gas
*Incomplete combustion of boiler fuel
*Radiant losses from the boiler exterior
*Blowdown losses
*Distribution losses (pipe transport, steam traps, etc.)

The first four of these losses take place at the boiler and contribute to the heating efficiencies seen in Table 1 for steam created with coal, gas, and oil. Several methods can be used to minimize these losses, one of the foremost being the control of air-to-fuel ratio in the boiler. This ratio must be managed by weighing losses due to uncombusted fuel against losses due to the heat content of excess exhaust gas. (Broughton, 1994)

[[File:Air-fuel ratio.PNG|thumb|border|center|middle|upright=4|link=|atl=|Air-to-fuel ratio relationship with heat losses (Broughton Fig 2.3)]]

The air-to-fuel ratio can be optimized using a feedback process controller. The control system will analyze the oxygen content of exhaust air and adjust the incoming air flow rate to achieve a set percentage of excess air. While desired excess oxygen will vary depending on the type of fuel, it is consistently seen that in the minimum loss-range a 1% increase in excess air will result in a 1% decrease in efficiency. ("Oxygen Control")

In addition to air-to-fuel ratio management, steam heat losses at the boiler can be mitigated via energy recovery, which is discussed in further detail [[Utility_systems#Energy_Recovery|below]].

Outside the boiler, losses in distribution of steam throughout facilities can decrease energy efficiency by a significant amount. This can cause up to a 60% increase in losses, but typically results in an overall steam efficiency of 50-55% (down from approximately 75% at the boiler). (Broughton, 1994) There are several ways that this issue can be addressed. First and foremost are steam metering systems, which can be used to monitor heat supply to separate sections of the process facility, check efficiency of fuel use, and determine which processes in a given plant have the highest steam requirements. Another method of minimizing transport losses is to decentralize steam generation systems. It can be advantageous to have numerous smaller boilers rather than a single centralized boiler so that steam does not have to travel as long of distances to reach its destination.

==Cold Utility Efficiency==
Efficiency in cooling processes depends more on the method used, and by extension the amount of coolant needed. Water and air utility efficiencies depend primarily on the fluid flow required to maintain the system at a desired temperature, while powered refrigeration utilities (for colder processes) have efficiencies at approximately 60%--but ranging up to 90%--of Carnot cycle efficiency, a metric of ideal refrigeration efficiency. (Towler and Sinnott, 2012) Cooling systems represent by definition a loss of energy from the main process to the utility stream, and as such it is often useful to find other uses for the heated media before discharge.

==Energy Recovery==
Recovery and recycle of energy is perhaps the most important aspect of creating an energy efficient plant design, and it is important for process engineers to fully consider possibilities for heat recovery in order to aid in economic viability.

===Process Heat Exchange===
Heat exchanger networks are a very common energy recovery method in industrial processes. These networks most frequently allow energy from heated product streams to be transferred to feed streams that must be brought up to process temperature. (Biegler, 1997) More information on the function and design of heat exchanger networks can be found on the [[Heat_exchanger#Heat_Exchanger_Networks|heat exchanger wiki page]]. The following are several examples of energy recovery via heat exchange that are used in industrial processes.

In distillation columns the bottoms and distillate effluents have the potential for energy exchange. Though the condenser at the top of the column cannot supply its waste heat to the reboiler due to their respective temperatures, the effluent streams can supply heat to the feed via a feed-effluent exchanger. This reduces the utility requirements to raise the feed to column temperature. (Biegler, 1997)

[[File:Feed_sterilization.jpg|thumb|border|right|middle|upright=4|link=|atl=|Feed sterilization schematic. (Towler and Sinnott Fig. 3.30)]]

Feed sterilization, commonly used in the food industry, is a common application for heat recovery through process stream heat exchange. In this application, the feed must be heated for a certain amount of time to kill any biological contaminants, after which it can be used to heat the new raw feed for sterilization. This reduces energy demands on the steam heater and thus reduces cost. (Towler and Sinnott, 2012)

In multi-vessel batch processes it can be advantageous to exchange heat as the process fluid is being transferred between vessels. Like the previous examples, this reduces the utility needed to bring the colder feed up to process temperature, thus reducing costs. (Towler and Sinnott, 2012)

[[File:Batch_heat_exchange.jpg|thumb|border|center|middle|upright=4|link=|atl=|Multi-vessel batch heat exchange schematic (Towler and Sinnott Fig. 3.31)]]

===Utility Regeneration===
[[File:Waste_heat_boiler.jpg|thumb|border|right|middle|200px|upright=4|link=|atl=|Industrial modular HRSG]]
When recovery of waste heat via transfer to other process streams is inconvenient or impossible, energy efficiency can still be improved through the regeneration of utilities. This is commonly done through the regeneration of steam by removing heat from exiting streams or from highly exothermic reactions. Waste heat in exiting streams can be removed via heat recovery steam generators (HRSGs), and is most often used on exiting gas streams. Heat recovery from reactions is a viable option when the reactor temperature will be at 150 C or above, as this will create steam at high enough pressure to be used in other processes. (Towler and Sinnott, 2012)

In the case of steam, waste heat and water treatment losses can be recovered from the utility generation process itself. One of the most common ways to do this is with an economizer. As seen in the schematic below, economizers have heat exchangers fit to the exhaust gas flow in order to transfer waste heat from these gases to the incoming boiler feed water. This can result in fuel energy savings of approximately 15% for typical excess air percentages in steam boilers. (Broughton, 1994; "Economizers")

[[File:Economizer.png|thumb|border|center|middle|upright=4|link=|atl=|Economizer schematic ("Economizers") and economizer energy recovery correlations (Broughton Fig 2.2)]]

[[File:Condensate recycle.PNG|thumb|border|right|middle|200px|upright=4|link=|atl=|Condensate return fuel savings (Broughton Fig 2.4)]]

Additionally, there are still more opportunities for recovery in steam generation processes. One typical case is the use of waste heat to preheat air entering the boiler. Optimistically, this can result in a 5% improvement in heat recovery for the system. Another method is the recycle of steam condensate back to the boiler. This has two main benefits: it retains the heat still present in the condensed steam and it retains the pretreatment chemicals added to the steam. (Broughton, 1994)

Furthermore, there is opportunity for energy recovery in the expansion of compressed gas through a turbine to create electricity, a process that can be economically viable given sufficiently high flows or pressure. One of the primary application of this type of energy recovery is in the creation of medium pressure and low pressure steam. In most processes, all steam is generated as high pressure steam and can be expanded through a turbine to decrease its pressure. Such technology has also been used in processes to synthesize ammonia, perform air separations, and synthesize nitric acid. (Towler and Sinnott, 2012) Recently, however, there has been a particularly strong interest for energy recovery in the natural gas industry, when gas is decompressed from major pipelines to residential low-pressure pipelines. A 2001 study estimated that there is the potential to recover 21 TWh, representing 11% of natural gas transport energy, via gas expansion. (Lehman)

=Process water and boiler-feed water=
Process water is water that will be directly used in the process. Boiler-feed water (BFW) is used to produce steam. Both may need to be purified to prevent impurities from contaminating a process or from foul equipment. It can be used as a cooling stream when the temperature of the stream to be cooled is greater than ~300 ⁰F. Cost of BFW can be partially offset by the steam credit.

Process water that undergoes moderate pretreatment can cost ~ $0.75/1,000 gal.

Extensive treatment ~ $6.00/1,000 gal.

Sterilized for pharmaceutical processes ~ $550/1,000 gal. (Seider pg 608)

=Demineralized Water=
In demineralized water, minerals have been removed by ion exchange. In boiler feed water, this is to prevent salt deposition, corrosion, formation of foam, and sluicing. In process water, the ions may contaminate the process.

=Waste Treatment=
Most chemical processes will produce some sort of waste. Disposal occurs to the atmosphere (in the case of some gases), sewers, body of water, or a landfill. Waste may require some treatment before disposal to meet regulations. Depending on process economics, byproducts may be recovered and processed. (Seider 2009 pg 609)

==Wastewater Treatment==
The United States EPA regulates industrial wastewater disposal through the Clean Water Act, introduced in 1948 as the Federal Water Pollution Control Act and amended to its current form in 1972. The sweeping 1972 amendments allowed the EPA to prevent industries and persons from discharging contaminated water into fresh water sources and set water quality standards. (Summary of the Clean Water Act) In accordance with this law, process plants in the United States treat wastewater at on-site or near-site treatment centers before releasing it into the surrounding environment.

Wastewater effluent streams, along with water runoff from around the plant, are treated to control for pH, toxicity, suspended solids, and biological oxygen demand (for aquatic life protection) prior to discharge. Each of these controls is typically addressed with a separate method. Acidity and basicity is balanced through the addition of an acid or alkaline solution. Toxic wastewater may be treated with chemical processes or simply diluted to safe concentrations. Suspended solids can be removed via filtration and/or with clarifiers. Oxygen demand of wastewater can be mitigated using activated sludge treatment processes. Once the water quality complies with the EPA, and state-mandated, regulations, it can be safely released. More information on the large number of industry-specific guidelines for waste effluent can be found on the EPA website (http://www.epa.gov/eg/industrial-effluent-guidelines).

=Air-Pollution Management=
===Introduction===

In the United States air pollution is regulated in the Clean Air Act, and almost all pollutant emitting plants are regulated under this law. The types of plants that can release significant emissions include petroleum refineries, sulfur recovery plants, carbon-black plants, fuel conversion plants, chemical process plants, fossil fuel plants, and petroleum storage and transfer facilities. To receive permission to construct a plant must undergo a review to show that it will not cause a violation of the Ambient Air Quality Standards(Peters, 2003).

==Methods==

There are two major types of pollutants that are released into the air, particulates and and gaseous pollutants. Particulates can be removed with mechanical forces while gaseous pollutants typically need to removed by chemical or physical means (Peters, 2003).

Particulates and volatile pollutants that need to be removed before disposal may be present. Particle removal equipment includes: cyclones, wet scrubbers, electrostatic precipitators, and fabric-filter systems (Seider 2009) such as bag filters (. The two charts below are from ''Plant Design and Economics for Chemical Engineers'' and show the types of equipment, separation methods, and particle sizes in different pollutant separation technologies (Peters, 1991).

[[File:AirPolChart.png|thumb|border|center|middle|upright=4|link=|atl=|Different Methods of Particle Separation, the Particle Sizes they Can Remove, and the Technologies Used]]

[[File:AirPolTable.png|thumb|border|center|middle|upright=4|link=|atl=| This Chart Shows the Specifications and Limitations of Different Separations Technologies Including Particle Size, Efficiencies, and Temperatures(Peters, 1991)]]

===Wet Scrubbers===
Wet scrubber use lime or limestone and water to remove SO2 and acid gases. The mixture can be injected into a scrubber or the gases can be bubbled through this mixture. This results in removal of 90-98% of SO2 and and acid gases (Clean Coal Technologies).

===Dry Scrubbers===
Dry scrubber blow powdered adsorbents into a vessel with gases and then after it has captured the SO2 and acid gases it is separated from gas using a fabric filter. These systems remove 90-93% of the contaminants (Clean Coal Technologies).

===Low-NOx Burners===
The purpose of low NOx burners is to decrease the amount of NOx created when the coal is burned. This is done by injecting coal and air in boilers. This can result in 40-50% NOx reduction. If air is injected into the area above the burner this can actually cause almost 70% NOx reduction (Clean Coal Technologies).

===Selective Non-Catalytic and Catalytic Reductions===
These systems inject ammonia into gases to remove NOx. The Catalytic reductions add a catalyst to the ammonia being injected to aid in the removal. Con-catalytic reductions result in about 35% removal, but adding a catalyst can increase that amount to about 90%. The catalytic-reduction can remove up to 80% of mercury as well (Clean Coal Technologies).

===Fabric Filters===
Also known as baghouses,these [[Separation processes#Filtration| filters]] remove particulates by passing air through filters. These can separate as much as 99.9% of particulate matter (Clean Coal Technologies).

===Electrostatic Precipitators===
Electrostatic Precicpitators remove particulate matter as gas passes through a device that has charged metal plates. The particles are then removed because of static electricity. These systems can remove between 99-99.9% of particulate matter (Clean Coal Technologies).

===Super-critical Boilers===
Super-Critical Boilers and Ultra-Supercritical Boilers operate at temperature and pressure higher than regular boilers. By operating at higher temperature these systems become more efficient. Super Critical Boilers typically have 10%-20% CO2 emissions than other similar sub-critical technologies. Ultra-Supercritical boilers can be as much as 30% more efficent than sub-critical technologies (Clean Coal Technologies).

===More Information===

For more information regarding some of these separations equipment see [[Solids-involved equipment]]

For more information Regarding Cyclones see [[Separation processes#Cyclones| Cyclones]], and for modeling cyclones see their[[Solids-involved equipment#HYSYS Simulation|HYSYS Simulation]]

Methods for removing inorganic and organic gaseous pollutants include: absorption, adsorption, condensation, and combustion (Seider 2009 pg 609). A list of typical of gases pollutants and their sources from ''Plant Design and Economics for Chemical Engineers'' is shown below (Peters, 1991).

[[File:GaseousPollutants.png|thumb|border|center|middle|upright=4|link=|atl=| Common Gaseous Pollutants and their Sources]]

Typically Gas-liquid absorption processes are done completed in a vertical, countercurrent, flow through packed, plate, or spray towers. These systems require good liquid-gas contact and proper equipment. These systems also often have significant energy consumption because of large pressure drops (Peters, 2003). For high volume systems absorption by scrubbing with water or another solvent is the most widely used method (Towler, 2012). Dry adsorbents can be used to remove the last races of gaseous pollutants. Adsorption typically requires blowers, condensers, separators, and controls. You also typically need two packed beds so that one can be used while the other is regenerated. Examples of adsorbents are molecular sieves and activated carbon. Incineration is typically used when there are gas streams that have no recovery value. This can be done with direct flame or catalytic oxidation. Catalytic oxidation usually has higher capital costs, but lower operating costs because it does not require fuel.

You can find more information on [[Separation processes#Absorption| Absorption]] and [[Separation processes#Adsorption| Adsorption]] in [[Separation processes]]

===Outcomes===

====United States Common Pollutant Emissions====

The United States implemented the Clean Air Act in 1970 and since then emissions in the U.S. have been drastically reduced. As a matter of fact despite increases in the population in the last 35 years the amount of pollutants emitted have decreased by almost 70% (EPA). More importantly this demonstrates that reasonable efforts can be put towards environmental protection without causing too much harm to industry.

[[File:EPA.png|thumb|border|center|middle|upright=4|link=|atl=| GDP and Other Growth Factors Vs Common Pollutants Released]]

===China Clean Coal===

====Success Through 2005====

Coal is a very inexpensive and abundant source of energy and is abundant in China (Xu 2010). In China Coal the cause of 90% of SO2 emissions, 70% of dust emissions, and 67% of NOx emissions, and 70% of CO2 emissions. These numbers are staggering considering that Coal has never been more than 50 percent of China's Energy Supply (Xu 2010).

[[File:Coal.jpg|thumb|border|center|middle|upright=4|link=|atl=| Coal Consumption for Power and the Percentage of Coal Used (Xu 2010)]]

Despite the increasing coal consumption, high efficiency electric dust removal systems with efficiencies that are as high as 99.6% have greatly decreased soot emissions to 32% below 1980's levels as of 2005.

[[File:Soot.jpg|thumb|border|center|middle|upright=4|link=|atl=| SO2 and Soot Emissions from 1981 to 2005 (Xu 2010)]]

====Technologies and Life Cycle Comparison====
As of 2010 China was consuming 48.2% of coal globally. Four potential technologies integrated gasification combined cycle (IGCC), sub-critical coal power generation (Sub-C); super-critical coal power generation (Super-C) ultrasuper-critical coal power generation (USC). These technologies are compared mostly on the basis of net generating efficiency and efficiency. Net generating efficiency is the output of the plant divided by the total available energy in the fuel used (Liang, 2013).

where efficiency is define as: η=E/Etotal

and

Etotal=Emining+Etransportation+Egeneration

[[File:technologies.PNG|thumb|border|center|middle|upright=4|link=|atl=|A comparison of many of the parameters important to lfe-cycle analysis (Liang, 2013)]]

This demonstrates that USC has the best net generation efficiency, while also having the largest capacity for a single system.

[[File:Lifecycle2.PNG|thumb|border|center|middle|upright=4|link=|atl=| (Liang, 2013)]]

This breaks down the energy efficiency of each system and demonstrates that USC is the most efficient

[[File:Lifecycle.PNG|thumb|border|center|middle|upright=4|link=|atl=| (Liang, 2013)]]

This demonstrates that the capital cost per unit of energy production is very high for IGCC, but shows that Super-C and USC are very competitive.

[[File:Price.PNG|thumb|border|center|middle|upright=4|link=|atl=| Capital cost in $/kW of clean coal power-generation technologies (Liang, 2013)]]

Lastly IGCC and USC are lowest on different element, but since these all have different global warming potential it is difficult to tell which is the most efficient.

And the table below defines all the relevant terms.

[[File:GWP.PNG|thumb|border|center|middle|upright=4|link=|atl=| Environemntal measures and their explanations (Liang, 2013)]]

All of the data suggests that USC is the highest in both energy efficiency and net generating efficiency while competitive in price and the lowest on emissions among the different boilers.

==Solid Waste==
U.S. federal regulations require that solid waste be classified as hazardous or nonhazardous. Conditions for a classification of hazardous include: ignitability, corrosivity, reactivity, toxicity, or posing a substantial threat to the surrounding environment and its inhabitants. Hazardous waste must be treated on- or near-site before being removed in containers. Non-hazardous waste may be landfilled or incinerated in some cases. A typical estimate of costs for waste disposal is $0.03/lb for nonhazardous solids and $0.10/lb for hazardous solids. (Seider 2009 pg 609)

=References=

*Biegler LT, Grossmann IE, Westerberg AW. ''Systematic Methods of Chemical Process Design'' Prentice-Hall: Upper Saddle River, 1997.

*Broughton, J. ''Process Utility Systems: Introduction to Design, Operation and Maintenance'' Institution of Chemical Engineers: Rugby, Warwickshire, UK, 1994.

*Clean Coal Technologies. America's Coalition for Clean Coal Electricity. http://www.americaspower.org/clean-coal-technologies-1663/

*Duke Energy Company (2013). How Do Coal Fired Plants Work? Charlotte: Duke Energy.

*"Economizers." Energy Solutions Center Inc.: Boiler Consortium. http://www.cleanboiler.org/Eff_Improve/Efficiency/Economizers.asp. Accessed February 21, 2016.

*Lehman B, Worrell E. ''Electricity Production from Natural Gas Pressure Recovery Using Expansion Turbines.'' Lawrence Berkeley National Laboratory; 2001.

*"Overview of the Clean Air Act and Air Pollution." Environmental Protection Agency. November 17, 2015. http://www.epa.gov/clean-air-act-overview. Accessed February 5, 2016.

*"Oxygen Control." Energy Solutions Center Inc.: Boiler Consortium. http://www.cleanboiler.org/Eff_Improve/Efficiency/Oxygen_Control.asp. Accessed February 21, 2016.

*Peters, Max S.; Timmerhaus, Klaus D.; West, Ronald E. (2003). " Plant Design and Economics for Chemical Engineers." McGraw Hill Higher Education.

*Ruina Xu Energy (2010). "Clean coal technology development in China." Elsevier.

*Seider, Seader, Lewin, Widagdo. (2009). ''Plant Design and Economics for Chemical Engineers, 5th Edition.'' Hoboken: Wiley.

*Seider, Seader, Lewin. (2008). ''Product and Process Design Principles, 2nd Edition.'' Hoboken: Wiley.

*Summary of the Clean Water Act. United States EPA website. http://www.epa.gov/laws-regulations/summary-clean-water-act

*Towler, G.P. and Sinnot, R. (2012). ''Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design.'' Elsevier.

*Turton R.; Bailie, R.C.; Whiting, W.B.; Shaeiwitz J.A.; Bhattacharyya D. (2012). ''Analysis, Synthesis, and Design of Chemical Processes.'' Upper Saddle River: Prentice Hall.

*G.D. Ulrich, A Guide to Chemical Engineering Process Design and Economics, Wiley: New York, 1984.

*Xiaoye Liang, Zhihua Wang, Zhijun Zhou, , Zhenyu Huang, Junhu Zhou, and Kefa Cen. (2013). "Up-to-date life cycle assessment and comparison study of clean coal power generation technologies in China" Elsevier.

Utility systems

2016-02-22T05:00:56Z

Bts511: /* Methods */

Authors: David Chen, [2014] Joshua Lee, [2015] Brett Sleyster, [2016] and Tom Aunins [2016] 

Stewards: David Chen, Jian Gong, and Fengqi You

Date Presented: January 13, 2014 /Date Revised: January 14, 2014

 


__TOC__

=Introduction=
Many chemical processes do not take place at ambient temperature or pressures. In order to reach these non-ambient conditions, utilities will have to be used to raise or lower temperatures and compress gases. (Towler, Towler/UOP) Utilities often contribute 5 to 10% of the price of a product, and may come from public or private utility companies or on-site plants. For purchased utilities, costs depend on consumption, while for company-owned utilities, there will be both capital and operating costs. They include things such as steam for heating, electricity, cooling water, refrigeration, fuels such as natural gas, wastewater treatment, waste disposal, and landfill. Steam is often the largest utility cost. Cogeneration unit can supply electricity accompanied with different steam pressures. (Seider 2009)

=Electricity=

Electricity is used to power many different kinds of equipment. It has many advantages: it is efficient (> 90%), reliable, available in a wide range of power, shaft speeds, designs, lifetimes, convenience, costs, and maintenance. It is generally used up to 200 hp, and sometimes over 10,000 Hp. In chemical process plants, the electricity demand is generally determined by the work or energy required for compression, pumping, air cooling, lights, and many other items. This electricity often times is purchased from local electricity providers, but many plants generate their own electricity via sophisticated processes.

Electricity is rarely used as a primary heat utility in large-scale chemical plants for a variety of reasons. The main disadvantages of using electricity as a heat utility are as follows (Towler 2012):

:*Heat from electricity is two to three times more expensive than heat from fuels. This is attributed to the lack of efficiency when creating heat from electricity.

:*Electrical heating units are expensive, require high maintenance, and must comply with strict safety regulations.

:*Electrical heating units are unsafe compared to steam heating units. In steam systems, the physically steam controls the temperature, whereas in electrical heating units temperature is controlled by temperature controllers, which can fail or burn out.

The use of electricity carries with it some hazards depending on the environment. Extra care must be taken when using electrically-powered equipment in areas which may have combustible fluids, vapors, or dust, and where liquids may be present. (Seider 2009 pg 606)

==Conventional Power Station==

In general, most electricity is generated from a conventional coal-fired process, whether it be on-site or purchased from a provider. Coal-fired processes have been used to create electricity throughout history, and technological advances have increased its efficiency and use worldwide. In a coal-fired steam station—much like a nuclear station—water is turned into steam, which in turn drives turbine generators to produce electricity. There are several variations on how to create energy from coal, but here are the basics of how a coal-fired process works:

[[File:coalfired.png|thumb|border|center|middle|upright=4|link=|atl=|General Coal-Fired Process Diagram (Duke Energy Company)]]

*Heat is created:
::Before the coal is burned, it is pulverized to the fineness of talcum powder. It is then mixed with hot air and blown into the firebox of the boiler. Burning in suspension, the coal/air mixture provides the most complete combustion and maximum heat possible.

*Water turns to steam:
::Highly purified water, pumped through pipes inside the boiler, is turned into steam by the heat. The steam reaches temperatures of up to 1,000 degrees Fahrenheit and pressures up to 3,500 pounds per square inch, and is piped to the turbine.

*Steam turns the turbine:
::The enormous pressure of the steam pushing against a series of giant turbine blades turns the turbine shaft. The turbine shaft is connected to the shaft of the generator, where magnets spin within wire coils to produce electricity.

*Steam is converted back to water:
::After doing its work in the turbine, the steam is drawn into a condenser, a large chamber in the basement of the power plant. In this important step, millions of gallons of cool water from a nearby source (such as a river or lake) are pumped through a network of tubes running through the condenser. The cool water in the tubes converts the steam back into water that can be used over and over again in the plant.

*Repeat:
::The cooling water is returned to its source without any contamination, and the steam water is returned to the boiler to repeat the cycle.

===Advantages & Disadvantages of Coal-Fired Energy Production===

[[File:coalfiredtable.png|thumb|border|center|middle|upright=4|link=|atl=|Advantages and Disadvantages of Coal-Fired Processes (Seider 2008)]]

==Gas-Turbine Cogeneration Process==

When generating energy on-site, many plants use a gas-turbine cogeneration process. The thermal efficiency of a gas-turbine process is in the range of 70-80% while conventional power stations, such as coal-fired processes, have a 30-40% efficiency. The lower efficiency in more conventional power stations is attributed to wasted heat in the exhaust steam in the condenser. One example of a gas-turbine process is outlined in the following figure. Figure 3.1 is a gas-turbine cogeneration process with a heat recovery steam generator (waste-heat) boiler.

[[File:cogeneration.png|thumb|border|center|middle|upright=4|link=|atl=|Gas-Turbine Cogeneration with a heat recovery steam generator boiler (Towler 2012 Fig 3.1)]]

Overall, the process illustrated is not much different from a coal-fired process. The main differences are that the cogeneration process creates both electricity and a heat utility, and cogeneration processes use natural gas instead of coal. Many of the advantages and disadvantages are similar to those of the coal-fired process, but the cogeneration has a much higher efficiency, creates heat to be used in another process, and uses a more volatile and expensive fuel. The main advantage of cogeneration over coal-fired energy production is that heat is not wasted. In coal-fired processes, heat is released and wasted during electricity generation. Cogeneration captures some, if not all of the byproduct for heat, which is an extremely useful utility that will be discussed in subsequent sections. In summary, the cogeneration plant is superior to the coal-fired process because of its higher efficiency and ability to create a useful heat utility.

Obviously any engineer would design the cogeneration plant to meet at least the energy requirement necessary for plant operation, but cogeneration plants often times are designed to exceed the plant electricity requirement to drive another source of capital. Many describe this scenario as a "make or buy" scenario (Towler 2012). This scenario provides chemical producers leverage when negotiating contracts with outsourced electricity providers and this allows plants to purchase electricity at a wholesale price. This is a huge advantage of considering on-site electricity production because electricity is needed in relatively high quantities for all chemical plants. Being able to minimize electricity costs, or even profit off of electricity production is a huge economical consideration that all plants employ.

=Process Heating=
The key objective of this section is to discuss how processes are heated. Heating utilities are necessary for proper usage of distillers, reactors, condensers, and several other integral types of equipment. More specifically, steam, fired heat, and hot oil/specialized heat transfer fluids will be discussed in the following subsections.

==Steam==
Steam is the most commonly used heat utility used in chemical plants, and as a result understanding how it is used is essential in the study of Utility systems. Steam is used both as a process fluid (feedstock, diluent to absorb heat of reaction, heating agent, and stripping agent in absorbers and adsorbers ) and utility. It can be used to drive pumps and compressors, ejectors (for producing a vacuum), and heat exchangers. As one can clearly see, steam is a versatile, and useful utility.

Here are a few advantages of using steam as opposed to other methods of process heating (Towler 2012):

:*By controlling the pressure of the steam, one can control the temperature at which the heat is released. Having a strong control over the temperature is essential in several processes. =
:*Steam is an efficient heat source because the heat of condensation of steam is very high. Meaning that there is is high output per mass of utility at a constant temperature.
:*Heat exchangers that use steam are relatively cheap because condensing steam has a high heat transfer coefficient.
:*Steam is nonflammable, nontoxic, and inert to several process fluids.

Chemical plants generally have a network of pipelines exclusively for providing steam. These networks generally have steam at a low pressure, a medium pressure, and a high pressure. The image below illustrates a basic steam system.

[[File:Steam.png|thumb|border|center|middle|upright=4|link=|atl=|(Towler 2012 Fig 3.2)]]

In the diagram above, boiler feed water at a high pressure is preheated and fed to other boilers. These other boilers superheat the steam to create a high pressure and high temperature steam stream. The steam is superheated past the dew point to account for heat loss in the pipelines. A portion of the high pressure steam is used for process heating in areas of the plant that require high temperatures. The rest of the high pressure steam is turned into medium pressure steam by valves and steam turbines. The medium pressure steam is then used to heat medium temperature processes and to form low pressure steam. The low pressure steam can be used to heat low pressure processes and it can be expanded in condensing turbines to create shaft work and energy. In summary, steam can be used for an innumerable amount of action items in a plant. High pressure, medium pressure, and low pressure steam can all be used as a heat source. Low pressure steam has utilities in creating electricity and it also has several other uses.

==Fired Heat==

In many cases, processes in a plant require a heat source stronger than high pressure and temperature steam. That is when fired heat is used, which is generally at temperatures above 523K. Streams can be heated directly in the furnace tubes or via a hot oil circuit or heat transfer fluid, which will be discussed in detail in the next section. Most fired heaters use natural gas as fuel because it burns cleaner than fuel oil. A cleaner burning fuel is always advantageous due to environmental and safety concerns. Furthermore, natural gases usually result in less wear and tear in burners and fuel lines.

Depending on the application of the fired-heater, different design specifications can be implemented to make the fired-heater as efficient as possible. The basic construction of a fired heater starts with a cylindrical chamber that is lined with with refractory bricks. The stream that is to be heated flows through tubes inside of the furnace. These tubes can be arranged in several different arrangements such as, around the walls of the furnace, or in horizontal or vertical banks (Towler 2012). The figure below illustrates the basic construction of the fired-heater and varying tube arrangements.

[[File:firedheater.png|thumb|border|center|middle|upright=4|link=|atl=|Basic Fired-Heater Diagrams (Towler 2012 Figure 19.66)]]

Fuel is burned to heat the entire furnace, and the heat transfer occurs from the combustion gases inside of the furnace across the tubes that are filled with our desired stream. The heat transfer between the tubes and the furnace is accomplished primarily via radiation. Modern designs take advantage of convective heat transfer by adding a smaller section on top of the fired-heater where the combustion gases flow over banks of tubes as seen in (c) in the figure above. Heat transfer can be further improved via convection by adding fins or pins in the combustion section.

The cost of fired heat can be calculated by the cost of fuel fired. Natural gas and heating oil are traded as commodities and prices can be found at many online trading sites or business news sites (i.e., www.cnn.money.com). Past historic prices for forecasting can be found in the Oil and Gas Journal or from the U.S. Energy Information Adminstration (www.eia.gov).

===Fuels===
Fuel is burned in utility facilities such as boilers, electricity generation, and cogeneration, and can be in solid, liquid, or gas form. It can also be burned to provide heating for a process or stream or to drive pumps and compressors. The fuel is usually burned with excess air to ensure complete combustion.

A way of quantifying the amount of heat generated is by using the heating values. Higher heating value (HHV) and the lower heating value (LHV) are used. The heating is the total heat evolved by complete combustion of a fuel with dry air with both at 60 ⁰F and the flue gas after combustion brought back down to ⁰F. If the water vapor in the flue gas is not condensed, we obtain the LHV. If the water vapor is condensed, the value of heat evolved is a bit higher, and this is the HHV.
Heating values for solids and liquids are usually on a per-mass basis, and gases on a per-volume basis.
A specified amount of heating cannot be met with the amount of fuel calculated using only the HHV. There will be heat losses, the flue gas temperature will be greater than 60 ⁰F, and water in the flue gas will typically be vapor. (Seider 608)

==Hot Oil/Specialized Heat Transfer Fluids==

Specialized heat transfer fluids and hot oil circuits are used as heat sources when steam and fired heat is not appropriate. Specialized heat transfer fluids and hot oil circuits are extremely versatile in that they can be used in the temperature range of 323K to 673K. This range however is quite variable. For hot oils, the upper temperature limit is gauged based off of the thermal decomposition of the oil and coking/fouling of heat exchanger tubes.

Hot oil circuit systems are most commonly used when the plant has several small temperature heating requirements because it is more economically sound. Rather than having several fired heaters heat each small temperature requirement, it is much more economical to have one fired heater heat the hot oil and circulate that oil through each of the process to meet all of the heat needs. Hot oil systems are also generally favored over high pressure steam in processes that involve high pressure differentials between the process stream and high pressure steam. Hot oil systems are favored in this scenario because of safety concerns. If the steam were to leak, the pressure drop could cause serious safety issues.

Mineral oils are the most commonly used heat transfer fluids, and one prominent example is Dowtherm A. Dowtherm A is a combination of 26.5 wt% diphenyl in diphenyl oxide (Towler 2012) and is extremely thermally stable. These mineral oil systems generally require high flow rates.

=Process Cooling=

==Cooling Water==
Cooling water is used to cool and/or condense streams. Cooling water is usually circulated between process heat exchangers and a cooling tower. Water is cooled during downward motion by contact with air blown upwards, which can bring the water temperature to come within ~ 5 ⁰F of air’s wet-bulb temperature.Approximately 80% of the temperature reduction is due to evaporation of the cooling water and heat transfer to the surrounding air. Water can also be cooled in spray ponds and cooling ponds. Both work by providing high area for water to exchange heat with air.
Water in cooling towers is lost through drift and blowdown, and makeup is usually 1.5 to 3% of the circulating rate.
If a large natural body of water is nearby, it can be used as a source of cooling water and discharged downstream. This water is usually filtered to remove salts and impurities that may lead to fouling, but it is not treated.

==Refrigeration==
Cooling water can usually be used to cool a stream to ~ 100 ⁰F. Air can only cool to ~ 120 ⁰F. Air may be used in places where water is scarce or more costly to transport. To cool or condense streams below 100 ⁰F, a refrigerant is typically used. Chilled brine can also be used, but is less common.

Until 1995, CFC Freon R-12 (dichlorodifuloromethane) and HCFC Freon R-22 (chlorodifuloromethane) were commonly used refrigerants. However, the chlorine atom in the molecules caused the depletion of the ozone layer. The U.S. Clean Air Act Amendments of 1990 went into effect in 1995, and the production of these refrigerants has since stopped or been greatly reduced.

Cost estimates are based on ton-day of refrigeration, where a ton is the heat that needs to be removed to freeze 1 ton per day of water at 32 ⁰F. Substitutes have since been developed. R-134a is often used in place of R-12. According to the EPA, R-134a is not combustible at ambient conditions, and is essentially non-toxic under 400 ppm, and is not ozone-depleting. (Seider pg 607)

=Energy Efficiency=
One of the chief concerns in selecting and designing process utility systems for heating and cooling is how to achieve the most energy efficient design. There are countless means by which plants lose energy, two of the foremost being through the mixing of different temperature or pressure streams and through the disposal of warmed cooling water. (Seider, 2009) Proper utilities design can help mitigate each of these losses as well as many others. The energy efficiency of a plant will depend primarily on the heating and cooling methods that are being used and the overall system design itself. These two parameters are important in determining how well energy is transferred to the desired media as well as how well that energy is recovered and recycled.

==Hot Utility Efficiency==
As mentioned above, the most commonly used utilities for process heating in large scale processes are steam, fired heat, and hot oil heaters. Of these, steam is the most commonly used. Electricity, while efficient at creating power, is not a viable source of heat in large industrial processes. Common ranges of heating efficiency for these three methods are displayed in Table 1. (Towler and Sinnott, 2012; Broughton, 1994)

{| class="wikitable"
|+Table 1: Process Heating Efficiencies
|-
! Process Heating Method
! Typical Efficiency
|-
| Steam (via coal boiler)
| 72%
|-
| Steam (via gas boiler)
| 73%
|-
| Steam (via oil boiler)
| 75%
|-
| Fired Heat w/ Convective Section
| 85%
|-
| Fired Heat w/o Convective Section
| 60%
|-
| Hot Oil Heaters/Vaporizers
| 80-85%
|}

===Steam===
As steam is so popular for heating purposes, it is useful to analyze the numerous ways in which losses can occur in steam systems. There are five primary sources of inefficiency and heat loss in the generation and distribution of steam throughout a process plant:

*The heat content of boiler exhaust gas
*Incomplete combustion of boiler fuel
*Radiant losses from the boiler exterior
*Blowdown losses
*Distribution losses (pipe transport, steam traps, etc.)

The first four of these losses take place at the boiler and contribute to the heating efficiencies seen in Table 1 for steam created with coal, gas, and oil. Several methods can be used to minimize these losses, one of the foremost being the control of air-to-fuel ratio in the boiler. This ratio must be managed by weighing losses due to uncombusted fuel against losses due to the heat content of excess exhaust gas. (Broughton, 1994)

[[File:Air-fuel ratio.PNG|thumb|border|center|middle|upright=4|link=|atl=|Air-to-fuel ratio relationship with heat losses (Broughton Fig 2.3)]]

The air-to-fuel ratio can be optimized using a feedback process controller. The control system will analyze the oxygen content of exhaust air and adjust the incoming air flow rate to achieve a set percentage of excess air. While desired excess oxygen will vary depending on the type of fuel, it is consistently seen that in the minimum loss-range a 1% increase in excess air will result in a 1% decrease in efficiency. ("Oxygen Control")

In addition to air-to-fuel ratio management, steam heat losses at the boiler can be mitigated via energy recovery, which is discussed in further detail [[Utility_systems#Energy_Recovery|below]].

Outside the boiler, losses in distribution of steam throughout facilities can decrease energy efficiency by a significant amount. This can cause up to a 60% increase in losses, but typically results in an overall steam efficiency of 50-55% (down from approximately 75% at the boiler). (Broughton, 1994) There are several ways that this issue can be addressed. First and foremost are steam metering systems, which can be used to monitor heat supply to separate sections of the process facility, check efficiency of fuel use, and determine which processes in a given plant have the highest steam requirements. Another method of minimizing transport losses is to decentralize steam generation systems. It can be advantageous to have numerous smaller boilers rather than a single centralized boiler so that steam does not have to travel as long of distances to reach its destination.

==Cold Utility Efficiency==
Efficiency in cooling processes depends more on the method used, and by extension the amount of coolant needed. Water and air utility efficiencies depend primarily on the fluid flow required to maintain the system at a desired temperature, while powered refrigeration utilities (for colder processes) have efficiencies at approximately 60%--but ranging up to 90%--of Carnot cycle efficiency, a metric of ideal refrigeration efficiency. (Towler and Sinnott, 2012) Cooling systems represent by definition a loss of energy from the main process to the utility stream, and as such it is often useful to find other uses for the heated media before discharge.

==Energy Recovery==
Recovery and recycle of energy is perhaps the most important aspect of creating an energy efficient plant design, and it is important for process engineers to fully consider possibilities for heat recovery in order to aid in economic viability.

===Process Heat Exchange===
Heat exchanger networks are a very common energy recovery method in industrial processes. These networks most frequently allow energy from heated product streams to be transferred to feed streams that must be brought up to process temperature. (Biegler, 1997) More information on the function and design of heat exchanger networks can be found on the [[Heat_exchanger#Heat_Exchanger_Networks|heat exchanger wiki page]]. The following are several examples of energy recovery via heat exchange that are used in industrial processes.

In distillation columns the bottoms and distillate effluents have the potential for energy exchange. Though the condenser at the top of the column cannot supply its waste heat to the reboiler due to their respective temperatures, the effluent streams can supply heat to the feed via a feed-effluent exchanger. This reduces the utility requirements to raise the feed to column temperature. (Biegler, 1997)

[[File:Feed_sterilization.jpg|thumb|border|right|middle|upright=4|link=|atl=|Feed sterilization schematic. (Towler and Sinnott Fig. 3.30)]]

Feed sterilization, commonly used in the food industry, is a common application for heat recovery through process stream heat exchange. In this application, the feed must be heated for a certain amount of time to kill any biological contaminants, after which it can be used to heat the new raw feed for sterilization. This reduces energy demands on the steam heater and thus reduces cost. (Towler and Sinnott, 2012)

In multi-vessel batch processes it can be advantageous to exchange heat as the process fluid is being transferred between vessels. Like the previous examples, this reduces the utility needed to bring the colder feed up to process temperature, thus reducing costs. (Towler and Sinnott, 2012)

[[File:Batch_heat_exchange.jpg|thumb|border|center|middle|upright=4|link=|atl=|Multi-vessel batch heat exchange schematic (Towler and Sinnott Fig. 3.31)]]

===Utility Regeneration===
[[File:Waste_heat_boiler.jpg|thumb|border|right|middle|200px|upright=4|link=|atl=|Industrial modular HRSG]]
When recovery of waste heat via transfer to other process streams is inconvenient or impossible, energy efficiency can still be improved through the regeneration of utilities. This is commonly done through the regeneration of steam by removing heat from exiting streams or from highly exothermic reactions. Waste heat in exiting streams can be removed via heat recovery steam generators (HRSGs), and is most often used on exiting gas streams. Heat recovery from reactions is a viable option when the reactor temperature will be at 150 C or above, as this will create steam at high enough pressure to be used in other processes. (Towler and Sinnott, 2012)

In the case of steam, waste heat and water treatment losses can be recovered from the utility generation process itself. One of the most common ways to do this is with an economizer. As seen in the schematic below, economizers have heat exchangers fit to the exhaust gas flow in order to transfer waste heat from these gases to the incoming boiler feed water. This can result in fuel energy savings of approximately 15% for typical excess air percentages in steam boilers. (Broughton, 1994; "Economizers")

[[File:Economizer.png|thumb|border|center|middle|upright=4|link=|atl=|Economizer schematic ("Economizers") and economizer energy recovery correlations (Broughton Fig 2.2)]]

[[File:Condensate recycle.PNG|thumb|border|right|middle|200px|upright=4|link=|atl=|Condensate return fuel savings (Broughton Fig 2.4)]]

Additionally, there are still more opportunities for recovery in steam generation processes. One typical case is the use of waste heat to preheat air entering the boiler. Optimistically, this can result in a 5% improvement in heat recovery for the system. Another method is the recycle of steam condensate back to the boiler. This has two main benefits: it retains the heat still present in the condensed steam and it retains the pretreatment chemicals added to the steam. (Broughton, 1994)

Furthermore, there is opportunity for energy recovery in the expansion of compressed gas through a turbine to create electricity, a process that can be economically viable given sufficiently high flows or pressure. One of the primary application of this type of energy recovery is in the creation of medium pressure and low pressure steam. In most processes, all steam is generated as high pressure steam and can be expanded through a turbine to decrease its pressure. Such technology has also been used in processes to synthesize ammonia, perform air separations, and synthesize nitric acid. (Towler and Sinnott, 2012) Recently, however, there has been a particularly strong interest for energy recovery in the natural gas industry, when gas is decompressed from major pipelines to residential low-pressure pipelines. A 2001 study estimated that there is the potential to recover 21 TWh, representing 11% of natural gas transport energy, via gas expansion. (Lehman)

=Process water and boiler-feed water=
Process water is water that will be directly used in the process. Boiler-feed water (BFW) is used to produce steam. Both may need to be purified to prevent impurities from contaminating a process or from foul equipment. It can be used as a cooling stream when the temperature of the stream to be cooled is greater than ~300 ⁰F. Cost of BFW can be partially offset by the steam credit.

Process water that undergoes moderate pretreatment can cost ~ $0.75/1,000 gal.

Extensive treatment ~ $6.00/1,000 gal.

Sterilized for pharmaceutical processes ~ $550/1,000 gal. (Seider pg 608)

=Demineralized Water=
In demineralized water, minerals have been removed by ion exchange. In boiler feed water, this is to prevent salt deposition, corrosion, formation of foam, and sluicing. In process water, the ions may contaminate the process.

=Waste Treatment=
Most chemical processes will produce some sort of waste. Disposal occurs to the atmosphere (in the case of some gases), sewers, body of water, or a landfill. Waste may require some treatment before disposal to meet regulations. Depending on process economics, byproducts may be recovered and processed. (Seider 2009 pg 609)

==Wastewater Treatment==
The United States EPA regulates industrial wastewater disposal through the Clean Water Act, introduced in 1948 as the Federal Water Pollution Control Act and amended to its current form in 1972. The sweeping 1972 amendments allowed the EPA to prevent industries and persons from discharging contaminated water into fresh water sources and set water quality standards. (Summary of the Clean Water Act) In accordance with this law, process plants in the United States treat wastewater at on-site or near-site treatment centers before releasing it into the surrounding environment.

Wastewater effluent streams, along with water runoff from around the plant, are treated to control for pH, toxicity, suspended solids, and biological oxygen demand (for aquatic life protection) prior to discharge. Each of these controls is typically addressed with a separate method. Acidity and basicity is balanced through the addition of an acid or alkaline solution. Toxic wastewater may be treated with chemical processes or simply diluted to safe concentrations. Suspended solids can be removed via filtration and/or with clarifiers. Oxygen demand of wastewater can be mitigated using activated sludge treatment processes. Once the water quality complies with the EPA, and state-mandated, regulations, it can be safely released. More information on the large number of industry-specific guidelines for waste effluent can be found on the EPA website (http://www.epa.gov/eg/industrial-effluent-guidelines).

=Air-Pollution Management=
===Introduction===

In the United States air pollution is regulated in the Clean Air Act, and almost all pollutant emitting plants are regulated under this law. The types of plants that can release significant emissions include petroleum refineries, sulfur recovery plants, carbon-black plants, fuel conversion plants, chemical process plants, fossil fuel plants, and petroleum storage and transfer facilities. To receive permission to construct a plant must undergo a review to show that it will not cause a violation of the Ambient Air Quality Standards(Peters, 2003).

==Methods==

There are two major types of pollutants that are released into the air, particulates and and gaseous pollutants. Particulates can be removed with mechanical forces while gaseous pollutants typically need to removed by chemical or physical means (Peters, 2003).

Particulates and volatile pollutants that need to be removed before disposal may be present. Particle removal equipment includes: cyclones, wet scrubbers, electrostatic precipitators, and fabric-filter systems (Seider 2009) such as bag filters (. The two charts below are from ''Plant Design and Economics for Chemical Engineers'' and show the types of equipment, separation methods, and particle sizes in different pollutant separation technologies (Peters, 1991).

[[File:AirPolChart.png|thumb|border|center|middle|upright=4|link=|atl=|Different Methods of Particle Separation, the Particle Sizes they Can Remove, and the Technologies Used]]

[[File:AirPolTable.png|thumb|border|center|middle|upright=4|link=|atl=| This Chart Shows the Specifications and Limitations of Different Separations Technologies Including Particle Size, Efficiencies, and Temperatures(Peters, 1991)]]

==Wet Scrubbers==
Wet scrubber use lime or limestone and water to remove SO2 and acid gases. The mixture can be injected into a scrubber or the gases can be bubbled through this mixture. This results in removal of 90-98% of SO2 and and acid gases (Clean Coal Technologies).

===Dry Scrubbers===
Dry scrubber blow powdered adsorbents into a vessel with gases and then after it has captured the SO2 and acid gases it is separated from gas using a fabric filter. These systems remove 90-93% of the contaminants (Clean Coal Technologies).

===Low-NOx Burners===
The purpose of low NOx burners is to decrease the amount of NOx created when the coal is burned. This is done by injecting coal and air in boilers. This can result in 40-50% NOx reduction. If air is injected into the area above the burner this can actually cause almost 70% NOx reduction (Clean Coal Technologies).

===Selective Non-Catalytic and Catalytic Reductions===
These systems inject ammonia into gases to remove NOx. The Catalytic reductions add a catalyst to the ammonia being injected to aid in the removal. Con-catalytic reductions result in about 35% removal, but adding a catalyst can increase that amount to about 90%. The catalytic-reduction can remove up to 80% of mercury as well (Clean Coal Technologies).

===Fabric Filters===
Also known as baghouses,these [[Separation processes#Filtration| filters]] remove particulates by passing air through filters. These can separate as much as 99.9% of particulate matter (Clean Coal Technologies).

===Electrostatic Precipitators===
Electrostatic Precicpitators remove particulate matter as gas passes through a device that has charged metal plates. The particles are then removed because of static electricity. These systems can remove between 99-99.9% of particulate matter (Clean Coal Technologies).

====Super-critical Boilers====
Super-Critical Boilers and Ultra-Supercritical Boilers operate at temperature and pressure higher than regular boilers. By operating at higher temperature these systems become more efficient. Super Critical Boilers typically have 10%-20% CO2 emissions than other similar sub-critical technologies. Ultra-Supercritical boilers can be as much as 30% more efficent than sub-critical technologies (Clean Coal Technologies).

===More Information===

For more information regarding some of these separations equipment see [[Solids-involved equipment]]

For more information Regarding Cyclones see [[Separation processes#Cyclones| Cyclones]], and for modeling cyclones see their[[Solids-involved equipment#HYSYS Simulation|HYSYS Simulation]]

Methods for removing inorganic and organic gaseous pollutants include: absorption, adsorption, condensation, and combustion (Seider 2009 pg 609). A list of typical of gases pollutants and their sources from ''Plant Design and Economics for Chemical Engineers'' is shown below (Peters, 1991).

[[File:GaseousPollutants.png|thumb|border|center|middle|upright=4|link=|atl=| Common Gaseous Pollutants and their Sources]]

Typically Gas-liquid absorption processes are done completed in a vertical, countercurrent, flow through packed, plate, or spray towers. These systems require good liquid-gas contact and proper equipment. These systems also often have significant energy consumption because of large pressure drops (Peters, 2003). For high volume systems absorption by scrubbing with water or another solvent is the most widely used method (Towler, 2012). Dry adsorbents can be used to remove the last races of gaseous pollutants. Adsorption typically requires blowers, condensers, separators, and controls. You also typically need two packed beds so that one can be used while the other is regenerated. Examples of adsorbents are molecular sieves and activated carbon. Incineration is typically used when there are gas streams that have no recovery value. This can be done with direct flame or catalytic oxidation. Catalytic oxidation usually has higher capital costs, but lower operating costs because it does not require fuel.

You can find more information on [[Separation processes#Absorption| Absorption]] and [[Separation processes#Adsorption| Adsorption]] in [[Separation processes]]

===Outcomes===

====United States Common Pollutant Emissions====

The United States implemented the Clean Air Act in 1970 and since then emissions in the U.S. have been drastically reduced. As a matter of fact despite increases in the population in the last 35 years the amount of pollutants emitted have decreased by almost 70% (EPA). More importantly this demonstrates that reasonable efforts can be put towards environmental protection without causing too much harm to industry.

[[File:EPA.png|thumb|border|center|middle|upright=4|link=|atl=| GDP and Other Growth Factors Vs Common Pollutants Released]]

===China Clean Coal===

====Success Through 2005====

Coal is a very inexpensive and abundant source of energy and is abundant in China (Xu 2010). In China Coal the cause of 90% of SO2 emissions, 70% of dust emissions, and 67% of NOx emissions, and 70% of CO2 emissions. These numbers are staggering considering that Coal has never been more than 50 percent of China's Energy Supply (Xu 2010).

[[File:Coal.jpg|thumb|border|center|middle|upright=4|link=|atl=| Coal Consumption for Power and the Percentage of Coal Used (Xu 2010)]]

Despite the increasing coal consumption, high efficiency electric dust removal systems with efficiencies that are as high as 99.6% have greatly decreased soot emissions to 32% below 1980's levels as of 2005.

[[File:Soot.jpg|thumb|border|center|middle|upright=4|link=|atl=| SO2 and Soot Emissions from 1981 to 2005 (Xu 2010)]]

====Technologies and Life Cycle Comparison====
As of 2010 China was consuming 48.2% of coal globally. Four potential technologies integrated gasification combined cycle (IGCC), sub-critical coal power generation (Sub-C); super-critical coal power generation (Super-C) ultrasuper-critical coal power generation (USC). These technologies are compared mostly on the basis of net generating efficiency and efficiency. Net generating efficiency is the output of the plant divided by the total available energy in the fuel used (Liang, 2013).

where efficiency is define as: η=E/Etotal

and

Etotal=Emining+Etransportation+Egeneration

[[File:technologies.PNG|thumb|border|center|middle|upright=4|link=|atl=|A comparison of many of the parameters important to lfe-cycle analysis (Liang, 2013)]]

This demonstrates that USC has the best net generation efficiency, while also having the largest capacity for a single system.

[[File:Lifecycle2.PNG|thumb|border|center|middle|upright=4|link=|atl=| (Liang, 2013)]]

This breaks down the energy efficiency of each system and demonstrates that USC is the most efficient

[[File:Lifecycle.PNG|thumb|border|center|middle|upright=4|link=|atl=| (Liang, 2013)]]

This demonstrates that the capital cost per unit of energy production is very high for IGCC, but shows that Super-C and USC are very competitive.

[[File:Price.PNG|thumb|border|center|middle|upright=4|link=|atl=| Capital cost in $/kW of clean coal power-generation technologies (Liang, 2013)]]

Lastly IGCC and USC are lowest on different element, but since these all have different global warming potential it is difficult to tell which is the most efficient.

And the table below defines all the relevant terms.

[[File:GWP.PNG|thumb|border|center|middle|upright=4|link=|atl=| Environemntal measures and their explanations (Liang, 2013)]]

All of the data suggests that USC is the highest in both energy efficiency and net generating efficiency while competitive in price and the lowest on emissions among the different boilers.

==Solid Waste==
U.S. federal regulations require that solid waste be classified as hazardous or nonhazardous. Conditions for a classification of hazardous include: ignitability, corrosivity, reactivity, toxicity, or posing a substantial threat to the surrounding environment and its inhabitants. Hazardous waste must be treated on- or near-site before being removed in containers. Non-hazardous waste may be landfilled or incinerated in some cases. A typical estimate of costs for waste disposal is $0.03/lb for nonhazardous solids and $0.10/lb for hazardous solids. (Seider 2009 pg 609)

=References=

*Biegler LT, Grossmann IE, Westerberg AW. ''Systematic Methods of Chemical Process Design'' Prentice-Hall: Upper Saddle River, 1997.

*Broughton, J. ''Process Utility Systems: Introduction to Design, Operation and Maintenance'' Institution of Chemical Engineers: Rugby, Warwickshire, UK, 1994.

*Clean Coal Technologies. America's Coalition for Clean Coal Electricity. http://www.americaspower.org/clean-coal-technologies-1663/

*Duke Energy Company (2013). How Do Coal Fired Plants Work? Charlotte: Duke Energy.

*"Economizers." Energy Solutions Center Inc.: Boiler Consortium. http://www.cleanboiler.org/Eff_Improve/Efficiency/Economizers.asp. Accessed February 21, 2016.

*Lehman B, Worrell E. ''Electricity Production from Natural Gas Pressure Recovery Using Expansion Turbines.'' Lawrence Berkeley National Laboratory; 2001.

*"Overview of the Clean Air Act and Air Pollution." Environmental Protection Agency. November 17, 2015. http://www.epa.gov/clean-air-act-overview. Accessed February 5, 2016.

*"Oxygen Control." Energy Solutions Center Inc.: Boiler Consortium. http://www.cleanboiler.org/Eff_Improve/Efficiency/Oxygen_Control.asp. Accessed February 21, 2016.

*Peters, Max S.; Timmerhaus, Klaus D.; West, Ronald E. (2003). " Plant Design and Economics for Chemical Engineers." McGraw Hill Higher Education.

*Ruina Xu Energy (2010). "Clean coal technology development in China." Elsevier.

*Seider, Seader, Lewin, Widagdo. (2009). ''Plant Design and Economics for Chemical Engineers, 5th Edition.'' Hoboken: Wiley.

*Seider, Seader, Lewin. (2008). ''Product and Process Design Principles, 2nd Edition.'' Hoboken: Wiley.

*Summary of the Clean Water Act. United States EPA website. http://www.epa.gov/laws-regulations/summary-clean-water-act

*Towler, G.P. and Sinnot, R. (2012). ''Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design.'' Elsevier.

*Turton R.; Bailie, R.C.; Whiting, W.B.; Shaeiwitz J.A.; Bhattacharyya D. (2012). ''Analysis, Synthesis, and Design of Chemical Processes.'' Upper Saddle River: Prentice Hall.

*G.D. Ulrich, A Guide to Chemical Engineering Process Design and Economics, Wiley: New York, 1984.

*Xiaoye Liang, Zhihua Wang, Zhijun Zhou, , Zhenyu Huang, Junhu Zhou, and Kefa Cen. (2013). "Up-to-date life cycle assessment and comparison study of clean coal power generation technologies in China" Elsevier.

Utility systems

2016-02-22T04:57:30Z

Bts511: /* Air-Pollution Management */

Authors: David Chen, [2014] Joshua Lee, [2015] Brett Sleyster, [2016] and Tom Aunins [2016] 

Stewards: David Chen, Jian Gong, and Fengqi You

Date Presented: January 13, 2014 /Date Revised: January 14, 2014

 


__TOC__

=Introduction=
Many chemical processes do not take place at ambient temperature or pressures. In order to reach these non-ambient conditions, utilities will have to be used to raise or lower temperatures and compress gases. (Towler, Towler/UOP) Utilities often contribute 5 to 10% of the price of a product, and may come from public or private utility companies or on-site plants. For purchased utilities, costs depend on consumption, while for company-owned utilities, there will be both capital and operating costs. They include things such as steam for heating, electricity, cooling water, refrigeration, fuels such as natural gas, wastewater treatment, waste disposal, and landfill. Steam is often the largest utility cost. Cogeneration unit can supply electricity accompanied with different steam pressures. (Seider 2009)

=Electricity=

Electricity is used to power many different kinds of equipment. It has many advantages: it is efficient (> 90%), reliable, available in a wide range of power, shaft speeds, designs, lifetimes, convenience, costs, and maintenance. It is generally used up to 200 hp, and sometimes over 10,000 Hp. In chemical process plants, the electricity demand is generally determined by the work or energy required for compression, pumping, air cooling, lights, and many other items. This electricity often times is purchased from local electricity providers, but many plants generate their own electricity via sophisticated processes.

Electricity is rarely used as a primary heat utility in large-scale chemical plants for a variety of reasons. The main disadvantages of using electricity as a heat utility are as follows (Towler 2012):

:*Heat from electricity is two to three times more expensive than heat from fuels. This is attributed to the lack of efficiency when creating heat from electricity.

:*Electrical heating units are expensive, require high maintenance, and must comply with strict safety regulations.

:*Electrical heating units are unsafe compared to steam heating units. In steam systems, the physically steam controls the temperature, whereas in electrical heating units temperature is controlled by temperature controllers, which can fail or burn out.

The use of electricity carries with it some hazards depending on the environment. Extra care must be taken when using electrically-powered equipment in areas which may have combustible fluids, vapors, or dust, and where liquids may be present. (Seider 2009 pg 606)

==Conventional Power Station==

In general, most electricity is generated from a conventional coal-fired process, whether it be on-site or purchased from a provider. Coal-fired processes have been used to create electricity throughout history, and technological advances have increased its efficiency and use worldwide. In a coal-fired steam station—much like a nuclear station—water is turned into steam, which in turn drives turbine generators to produce electricity. There are several variations on how to create energy from coal, but here are the basics of how a coal-fired process works:

[[File:coalfired.png|thumb|border|center|middle|upright=4|link=|atl=|General Coal-Fired Process Diagram (Duke Energy Company)]]

*Heat is created:
::Before the coal is burned, it is pulverized to the fineness of talcum powder. It is then mixed with hot air and blown into the firebox of the boiler. Burning in suspension, the coal/air mixture provides the most complete combustion and maximum heat possible.

*Water turns to steam:
::Highly purified water, pumped through pipes inside the boiler, is turned into steam by the heat. The steam reaches temperatures of up to 1,000 degrees Fahrenheit and pressures up to 3,500 pounds per square inch, and is piped to the turbine.

*Steam turns the turbine:
::The enormous pressure of the steam pushing against a series of giant turbine blades turns the turbine shaft. The turbine shaft is connected to the shaft of the generator, where magnets spin within wire coils to produce electricity.

*Steam is converted back to water:
::After doing its work in the turbine, the steam is drawn into a condenser, a large chamber in the basement of the power plant. In this important step, millions of gallons of cool water from a nearby source (such as a river or lake) are pumped through a network of tubes running through the condenser. The cool water in the tubes converts the steam back into water that can be used over and over again in the plant.

*Repeat:
::The cooling water is returned to its source without any contamination, and the steam water is returned to the boiler to repeat the cycle.

===Advantages & Disadvantages of Coal-Fired Energy Production===

[[File:coalfiredtable.png|thumb|border|center|middle|upright=4|link=|atl=|Advantages and Disadvantages of Coal-Fired Processes (Seider 2008)]]

==Gas-Turbine Cogeneration Process==

When generating energy on-site, many plants use a gas-turbine cogeneration process. The thermal efficiency of a gas-turbine process is in the range of 70-80% while conventional power stations, such as coal-fired processes, have a 30-40% efficiency. The lower efficiency in more conventional power stations is attributed to wasted heat in the exhaust steam in the condenser. One example of a gas-turbine process is outlined in the following figure. Figure 3.1 is a gas-turbine cogeneration process with a heat recovery steam generator (waste-heat) boiler.

[[File:cogeneration.png|thumb|border|center|middle|upright=4|link=|atl=|Gas-Turbine Cogeneration with a heat recovery steam generator boiler (Towler 2012 Fig 3.1)]]

Overall, the process illustrated is not much different from a coal-fired process. The main differences are that the cogeneration process creates both electricity and a heat utility, and cogeneration processes use natural gas instead of coal. Many of the advantages and disadvantages are similar to those of the coal-fired process, but the cogeneration has a much higher efficiency, creates heat to be used in another process, and uses a more volatile and expensive fuel. The main advantage of cogeneration over coal-fired energy production is that heat is not wasted. In coal-fired processes, heat is released and wasted during electricity generation. Cogeneration captures some, if not all of the byproduct for heat, which is an extremely useful utility that will be discussed in subsequent sections. In summary, the cogeneration plant is superior to the coal-fired process because of its higher efficiency and ability to create a useful heat utility.

Obviously any engineer would design the cogeneration plant to meet at least the energy requirement necessary for plant operation, but cogeneration plants often times are designed to exceed the plant electricity requirement to drive another source of capital. Many describe this scenario as a "make or buy" scenario (Towler 2012). This scenario provides chemical producers leverage when negotiating contracts with outsourced electricity providers and this allows plants to purchase electricity at a wholesale price. This is a huge advantage of considering on-site electricity production because electricity is needed in relatively high quantities for all chemical plants. Being able to minimize electricity costs, or even profit off of electricity production is a huge economical consideration that all plants employ.

=Process Heating=
The key objective of this section is to discuss how processes are heated. Heating utilities are necessary for proper usage of distillers, reactors, condensers, and several other integral types of equipment. More specifically, steam, fired heat, and hot oil/specialized heat transfer fluids will be discussed in the following subsections.

==Steam==
Steam is the most commonly used heat utility used in chemical plants, and as a result understanding how it is used is essential in the study of Utility systems. Steam is used both as a process fluid (feedstock, diluent to absorb heat of reaction, heating agent, and stripping agent in absorbers and adsorbers ) and utility. It can be used to drive pumps and compressors, ejectors (for producing a vacuum), and heat exchangers. As one can clearly see, steam is a versatile, and useful utility.

Here are a few advantages of using steam as opposed to other methods of process heating (Towler 2012):

:*By controlling the pressure of the steam, one can control the temperature at which the heat is released. Having a strong control over the temperature is essential in several processes. =
:*Steam is an efficient heat source because the heat of condensation of steam is very high. Meaning that there is is high output per mass of utility at a constant temperature.
:*Heat exchangers that use steam are relatively cheap because condensing steam has a high heat transfer coefficient.
:*Steam is nonflammable, nontoxic, and inert to several process fluids.

Chemical plants generally have a network of pipelines exclusively for providing steam. These networks generally have steam at a low pressure, a medium pressure, and a high pressure. The image below illustrates a basic steam system.

[[File:Steam.png|thumb|border|center|middle|upright=4|link=|atl=|(Towler 2012 Fig 3.2)]]

In the diagram above, boiler feed water at a high pressure is preheated and fed to other boilers. These other boilers superheat the steam to create a high pressure and high temperature steam stream. The steam is superheated past the dew point to account for heat loss in the pipelines. A portion of the high pressure steam is used for process heating in areas of the plant that require high temperatures. The rest of the high pressure steam is turned into medium pressure steam by valves and steam turbines. The medium pressure steam is then used to heat medium temperature processes and to form low pressure steam. The low pressure steam can be used to heat low pressure processes and it can be expanded in condensing turbines to create shaft work and energy. In summary, steam can be used for an innumerable amount of action items in a plant. High pressure, medium pressure, and low pressure steam can all be used as a heat source. Low pressure steam has utilities in creating electricity and it also has several other uses.

==Fired Heat==

In many cases, processes in a plant require a heat source stronger than high pressure and temperature steam. That is when fired heat is used, which is generally at temperatures above 523K. Streams can be heated directly in the furnace tubes or via a hot oil circuit or heat transfer fluid, which will be discussed in detail in the next section. Most fired heaters use natural gas as fuel because it burns cleaner than fuel oil. A cleaner burning fuel is always advantageous due to environmental and safety concerns. Furthermore, natural gases usually result in less wear and tear in burners and fuel lines.

Depending on the application of the fired-heater, different design specifications can be implemented to make the fired-heater as efficient as possible. The basic construction of a fired heater starts with a cylindrical chamber that is lined with with refractory bricks. The stream that is to be heated flows through tubes inside of the furnace. These tubes can be arranged in several different arrangements such as, around the walls of the furnace, or in horizontal or vertical banks (Towler 2012). The figure below illustrates the basic construction of the fired-heater and varying tube arrangements.

[[File:firedheater.png|thumb|border|center|middle|upright=4|link=|atl=|Basic Fired-Heater Diagrams (Towler 2012 Figure 19.66)]]

Fuel is burned to heat the entire furnace, and the heat transfer occurs from the combustion gases inside of the furnace across the tubes that are filled with our desired stream. The heat transfer between the tubes and the furnace is accomplished primarily via radiation. Modern designs take advantage of convective heat transfer by adding a smaller section on top of the fired-heater where the combustion gases flow over banks of tubes as seen in (c) in the figure above. Heat transfer can be further improved via convection by adding fins or pins in the combustion section.

The cost of fired heat can be calculated by the cost of fuel fired. Natural gas and heating oil are traded as commodities and prices can be found at many online trading sites or business news sites (i.e., www.cnn.money.com). Past historic prices for forecasting can be found in the Oil and Gas Journal or from the U.S. Energy Information Adminstration (www.eia.gov).

===Fuels===
Fuel is burned in utility facilities such as boilers, electricity generation, and cogeneration, and can be in solid, liquid, or gas form. It can also be burned to provide heating for a process or stream or to drive pumps and compressors. The fuel is usually burned with excess air to ensure complete combustion.

A way of quantifying the amount of heat generated is by using the heating values. Higher heating value (HHV) and the lower heating value (LHV) are used. The heating is the total heat evolved by complete combustion of a fuel with dry air with both at 60 ⁰F and the flue gas after combustion brought back down to ⁰F. If the water vapor in the flue gas is not condensed, we obtain the LHV. If the water vapor is condensed, the value of heat evolved is a bit higher, and this is the HHV.
Heating values for solids and liquids are usually on a per-mass basis, and gases on a per-volume basis.
A specified amount of heating cannot be met with the amount of fuel calculated using only the HHV. There will be heat losses, the flue gas temperature will be greater than 60 ⁰F, and water in the flue gas will typically be vapor. (Seider 608)

==Hot Oil/Specialized Heat Transfer Fluids==

Specialized heat transfer fluids and hot oil circuits are used as heat sources when steam and fired heat is not appropriate. Specialized heat transfer fluids and hot oil circuits are extremely versatile in that they can be used in the temperature range of 323K to 673K. This range however is quite variable. For hot oils, the upper temperature limit is gauged based off of the thermal decomposition of the oil and coking/fouling of heat exchanger tubes.

Hot oil circuit systems are most commonly used when the plant has several small temperature heating requirements because it is more economically sound. Rather than having several fired heaters heat each small temperature requirement, it is much more economical to have one fired heater heat the hot oil and circulate that oil through each of the process to meet all of the heat needs. Hot oil systems are also generally favored over high pressure steam in processes that involve high pressure differentials between the process stream and high pressure steam. Hot oil systems are favored in this scenario because of safety concerns. If the steam were to leak, the pressure drop could cause serious safety issues.

Mineral oils are the most commonly used heat transfer fluids, and one prominent example is Dowtherm A. Dowtherm A is a combination of 26.5 wt% diphenyl in diphenyl oxide (Towler 2012) and is extremely thermally stable. These mineral oil systems generally require high flow rates.

=Process Cooling=

==Cooling Water==
Cooling water is used to cool and/or condense streams. Cooling water is usually circulated between process heat exchangers and a cooling tower. Water is cooled during downward motion by contact with air blown upwards, which can bring the water temperature to come within ~ 5 ⁰F of air’s wet-bulb temperature.Approximately 80% of the temperature reduction is due to evaporation of the cooling water and heat transfer to the surrounding air. Water can also be cooled in spray ponds and cooling ponds. Both work by providing high area for water to exchange heat with air.
Water in cooling towers is lost through drift and blowdown, and makeup is usually 1.5 to 3% of the circulating rate.
If a large natural body of water is nearby, it can be used as a source of cooling water and discharged downstream. This water is usually filtered to remove salts and impurities that may lead to fouling, but it is not treated.

==Refrigeration==
Cooling water can usually be used to cool a stream to ~ 100 ⁰F. Air can only cool to ~ 120 ⁰F. Air may be used in places where water is scarce or more costly to transport. To cool or condense streams below 100 ⁰F, a refrigerant is typically used. Chilled brine can also be used, but is less common.

Until 1995, CFC Freon R-12 (dichlorodifuloromethane) and HCFC Freon R-22 (chlorodifuloromethane) were commonly used refrigerants. However, the chlorine atom in the molecules caused the depletion of the ozone layer. The U.S. Clean Air Act Amendments of 1990 went into effect in 1995, and the production of these refrigerants has since stopped or been greatly reduced.

Cost estimates are based on ton-day of refrigeration, where a ton is the heat that needs to be removed to freeze 1 ton per day of water at 32 ⁰F. Substitutes have since been developed. R-134a is often used in place of R-12. According to the EPA, R-134a is not combustible at ambient conditions, and is essentially non-toxic under 400 ppm, and is not ozone-depleting. (Seider pg 607)

=Energy Efficiency=
One of the chief concerns in selecting and designing process utility systems for heating and cooling is how to achieve the most energy efficient design. There are countless means by which plants lose energy, two of the foremost being through the mixing of different temperature or pressure streams and through the disposal of warmed cooling water. (Seider, 2009) Proper utilities design can help mitigate each of these losses as well as many others. The energy efficiency of a plant will depend primarily on the heating and cooling methods that are being used and the overall system design itself. These two parameters are important in determining how well energy is transferred to the desired media as well as how well that energy is recovered and recycled.

==Hot Utility Efficiency==
As mentioned above, the most commonly used utilities for process heating in large scale processes are steam, fired heat, and hot oil heaters. Of these, steam is the most commonly used. Electricity, while efficient at creating power, is not a viable source of heat in large industrial processes. Common ranges of heating efficiency for these three methods are displayed in Table 1. (Towler and Sinnott, 2012; Broughton, 1994)

{| class="wikitable"
|+Table 1: Process Heating Efficiencies
|-
! Process Heating Method
! Typical Efficiency
|-
| Steam (via coal boiler)
| 72%
|-
| Steam (via gas boiler)
| 73%
|-
| Steam (via oil boiler)
| 75%
|-
| Fired Heat w/ Convective Section
| 85%
|-
| Fired Heat w/o Convective Section
| 60%
|-
| Hot Oil Heaters/Vaporizers
| 80-85%
|}

===Steam===
As steam is so popular for heating purposes, it is useful to analyze the numerous ways in which losses can occur in steam systems. There are five primary sources of inefficiency and heat loss in the generation and distribution of steam throughout a process plant:

*The heat content of boiler exhaust gas
*Incomplete combustion of boiler fuel
*Radiant losses from the boiler exterior
*Blowdown losses
*Distribution losses (pipe transport, steam traps, etc.)

The first four of these losses take place at the boiler and contribute to the heating efficiencies seen in Table 1 for steam created with coal, gas, and oil. Several methods can be used to minimize these losses, one of the foremost being the control of air-to-fuel ratio in the boiler. This ratio must be managed by weighing losses due to uncombusted fuel against losses due to the heat content of excess exhaust gas. (Broughton, 1994)

[[File:Air-fuel ratio.PNG|thumb|border|center|middle|upright=4|link=|atl=|Air-to-fuel ratio relationship with heat losses (Broughton Fig 2.3)]]

The air-to-fuel ratio can be optimized using a feedback process controller. The control system will analyze the oxygen content of exhaust air and adjust the incoming air flow rate to achieve a set percentage of excess air. While desired excess oxygen will vary depending on the type of fuel, it is consistently seen that in the minimum loss-range a 1% increase in excess air will result in a 1% decrease in efficiency. ("Oxygen Control")

In addition to air-to-fuel ratio management, steam heat losses at the boiler can be mitigated via energy recovery, which is discussed in further detail [[Utility_systems#Energy_Recovery|below]].

Outside the boiler, losses in distribution of steam throughout facilities can decrease energy efficiency by a significant amount. This can cause up to a 60% increase in losses, but typically results in an overall steam efficiency of 50-55% (down from approximately 75% at the boiler). (Broughton, 1994) There are several ways that this issue can be addressed. First and foremost are steam metering systems, which can be used to monitor heat supply to separate sections of the process facility, check efficiency of fuel use, and determine which processes in a given plant have the highest steam requirements. Another method of minimizing transport losses is to decentralize steam generation systems. It can be advantageous to have numerous smaller boilers rather than a single centralized boiler so that steam does not have to travel as long of distances to reach its destination.

==Cold Utility Efficiency==
Efficiency in cooling processes depends more on the method used, and by extension the amount of coolant needed. Water and air utility efficiencies depend primarily on the fluid flow required to maintain the system at a desired temperature, while powered refrigeration utilities (for colder processes) have efficiencies at approximately 60%--but ranging up to 90%--of Carnot cycle efficiency, a metric of ideal refrigeration efficiency. (Towler and Sinnott, 2012) Cooling systems represent by definition a loss of energy from the main process to the utility stream, and as such it is often useful to find other uses for the heated media before discharge.

==Energy Recovery==
Recovery and recycle of energy is perhaps the most important aspect of creating an energy efficient plant design, and it is important for process engineers to fully consider possibilities for heat recovery in order to aid in economic viability.

===Process Heat Exchange===
Heat exchanger networks are a very common energy recovery method in industrial processes. These networks most frequently allow energy from heated product streams to be transferred to feed streams that must be brought up to process temperature. (Biegler, 1997) More information on the function and design of heat exchanger networks can be found on the [[Heat_exchanger#Heat_Exchanger_Networks|heat exchanger wiki page]]. The following are several examples of energy recovery via heat exchange that are used in industrial processes.

In distillation columns the bottoms and distillate effluents have the potential for energy exchange. Though the condenser at the top of the column cannot supply its waste heat to the reboiler due to their respective temperatures, the effluent streams can supply heat to the feed via a feed-effluent exchanger. This reduces the utility requirements to raise the feed to column temperature. (Biegler, 1997)

[[File:Feed_sterilization.jpg|thumb|border|right|middle|upright=4|link=|atl=|Feed sterilization schematic. (Towler and Sinnott Fig. 3.30)]]

Feed sterilization, commonly used in the food industry, is a common application for heat recovery through process stream heat exchange. In this application, the feed must be heated for a certain amount of time to kill any biological contaminants, after which it can be used to heat the new raw feed for sterilization. This reduces energy demands on the steam heater and thus reduces cost. (Towler and Sinnott, 2012)

In multi-vessel batch processes it can be advantageous to exchange heat as the process fluid is being transferred between vessels. Like the previous examples, this reduces the utility needed to bring the colder feed up to process temperature, thus reducing costs. (Towler and Sinnott, 2012)

[[File:Batch_heat_exchange.jpg|thumb|border|center|middle|upright=4|link=|atl=|Multi-vessel batch heat exchange schematic (Towler and Sinnott Fig. 3.31)]]

===Utility Regeneration===
[[File:Waste_heat_boiler.jpg|thumb|border|right|middle|200px|upright=4|link=|atl=|Industrial modular HRSG]]
When recovery of waste heat via transfer to other process streams is inconvenient or impossible, energy efficiency can still be improved through the regeneration of utilities. This is commonly done through the regeneration of steam by removing heat from exiting streams or from highly exothermic reactions. Waste heat in exiting streams can be removed via heat recovery steam generators (HRSGs), and is most often used on exiting gas streams. Heat recovery from reactions is a viable option when the reactor temperature will be at 150 C or above, as this will create steam at high enough pressure to be used in other processes. (Towler and Sinnott, 2012)

In the case of steam, waste heat and water treatment losses can be recovered from the utility generation process itself. One of the most common ways to do this is with an economizer. As seen in the schematic below, economizers have heat exchangers fit to the exhaust gas flow in order to transfer waste heat from these gases to the incoming boiler feed water. This can result in fuel energy savings of approximately 15% for typical excess air percentages in steam boilers. (Broughton, 1994; "Economizers")

[[File:Economizer.png|thumb|border|center|middle|upright=4|link=|atl=|Economizer schematic ("Economizers") and economizer energy recovery correlations (Broughton Fig 2.2)]]

[[File:Condensate recycle.PNG|thumb|border|right|middle|200px|upright=4|link=|atl=|Condensate return fuel savings (Broughton Fig 2.4)]]

Additionally, there are still more opportunities for recovery in steam generation processes. One typical case is the use of waste heat to preheat air entering the boiler. Optimistically, this can result in a 5% improvement in heat recovery for the system. Another method is the recycle of steam condensate back to the boiler. This has two main benefits: it retains the heat still present in the condensed steam and it retains the pretreatment chemicals added to the steam. (Broughton, 1994)

Furthermore, there is opportunity for energy recovery in the expansion of compressed gas through a turbine to create electricity, a process that can be economically viable given sufficiently high flows or pressure. One of the primary application of this type of energy recovery is in the creation of medium pressure and low pressure steam. In most processes, all steam is generated as high pressure steam and can be expanded through a turbine to decrease its pressure. Such technology has also been used in processes to synthesize ammonia, perform air separations, and synthesize nitric acid. (Towler and Sinnott, 2012) Recently, however, there has been a particularly strong interest for energy recovery in the natural gas industry, when gas is decompressed from major pipelines to residential low-pressure pipelines. A 2001 study estimated that there is the potential to recover 21 TWh, representing 11% of natural gas transport energy, via gas expansion. (Lehman)

=Process water and boiler-feed water=
Process water is water that will be directly used in the process. Boiler-feed water (BFW) is used to produce steam. Both may need to be purified to prevent impurities from contaminating a process or from foul equipment. It can be used as a cooling stream when the temperature of the stream to be cooled is greater than ~300 ⁰F. Cost of BFW can be partially offset by the steam credit.

Process water that undergoes moderate pretreatment can cost ~ $0.75/1,000 gal.

Extensive treatment ~ $6.00/1,000 gal.

Sterilized for pharmaceutical processes ~ $550/1,000 gal. (Seider pg 608)

=Demineralized Water=
In demineralized water, minerals have been removed by ion exchange. In boiler feed water, this is to prevent salt deposition, corrosion, formation of foam, and sluicing. In process water, the ions may contaminate the process.

=Waste Treatment=
Most chemical processes will produce some sort of waste. Disposal occurs to the atmosphere (in the case of some gases), sewers, body of water, or a landfill. Waste may require some treatment before disposal to meet regulations. Depending on process economics, byproducts may be recovered and processed. (Seider 2009 pg 609)

==Wastewater Treatment==
The United States EPA regulates industrial wastewater disposal through the Clean Water Act, introduced in 1948 as the Federal Water Pollution Control Act and amended to its current form in 1972. The sweeping 1972 amendments allowed the EPA to prevent industries and persons from discharging contaminated water into fresh water sources and set water quality standards. (Summary of the Clean Water Act) In accordance with this law, process plants in the United States treat wastewater at on-site or near-site treatment centers before releasing it into the surrounding environment.

Wastewater effluent streams, along with water runoff from around the plant, are treated to control for pH, toxicity, suspended solids, and biological oxygen demand (for aquatic life protection) prior to discharge. Each of these controls is typically addressed with a separate method. Acidity and basicity is balanced through the addition of an acid or alkaline solution. Toxic wastewater may be treated with chemical processes or simply diluted to safe concentrations. Suspended solids can be removed via filtration and/or with clarifiers. Oxygen demand of wastewater can be mitigated using activated sludge treatment processes. Once the water quality complies with the EPA, and state-mandated, regulations, it can be safely released. More information on the large number of industry-specific guidelines for waste effluent can be found on the EPA website (http://www.epa.gov/eg/industrial-effluent-guidelines).

=Air-Pollution Management=
===Introduction===

In the United States air pollution is regulated in the Clean Air Act, and almost all pollutant emitting plants are regulated under this law. The types of plants that can release significant emissions include petroleum refineries, sulfur recovery plants, carbon-black plants, fuel conversion plants, chemical process plants, fossil fuel plants, and petroleum storage and transfer facilities. To receive permission to construct a plant must undergo a review to show that it will not cause a violation of the Ambient Air Quality Standards(Peters, 2003).

==Methods==

There are two major types of pollutants that are released into the air, particulates and and gaseous pollutants. Particulates can be removed with mechanical forces while gaseous pollutants typically need to removed by chemical or physical means (Peters, 2003).

Particulates and volatile pollutants that need to be removed before disposal may be present. Particle removal equipment includes: cyclones, wet scrubbers, electrostatic precipitators, and fabric-filter systems (Seider 2009) such as bag filters (. The two charts below are from ''Plant Design and Economics for Chemical Engineers'' and show the types of equipment, separation methods, and particle sizes in different pollutant separation technologies (Peters, 1991).

[[File:AirPolChart.png|thumb|border|center|middle|upright=4|link=|atl=|Different Methods of Particle Separation, the Particle Sizes they Can Remove, and the Technologies Used]]

[[File:AirPolTable.png|thumb|border|center|middle|upright=4|link=|atl=| This Chart Shows the Specifications and Limitations of Different Separations Technologies Including Particle Size, Efficiencies, and Temperatures(Peters, 1991)]]

====Wet Scrubbers====
Wet scrubber use lime or limestone and water to remove SO2 and acid gases. The mixture can be injected into a scrubber or the gases can be bubbled through this mixture. This results in removal of 90-98% of SO2 and and acid gases (Clean Coal Technologies).

====Dry Scrubbers====
Dry scrubber blow powdered adsorbents into a vessel with gases and then after it has captured the SO2 and acid gases it is separated from gas using a fabric filter. These systems remove 90-93% of the contaminants (Clean Coal Technologies).

====Low-NOx Burners====
The purpose of low NOx burners is to decrease the amount of NOx created when the coal is burned. This is done by injecting coal and air in boilers. This can result in 40-50% NOx reduction. If air is injected into the area above the burner this can actually cause almost 70% NOx reduction (Clean Coal Technologies).

====Selective Non-Catalytic and Catalytic Reductions====
These systems inject ammonia into gases to remove NOx. The Catalytic reductions add a catalyst to the ammonia being injected to aid in the removal. Con-catalytic reductions result in about 35% removal, but adding a catalyst can increase that amount to about 90%. The catalytic-reduction can remove up to 80% of mercury as well (Clean Coal Technologies).

====Fabric Filters====
Also known as baghouses,these [[Separation processes#Filtration| filters]] remove particulates by passing air through filters. These can separate as much as 99.9% of particulate matter (Clean Coal Technologies).

====Electrostatic Precipitators====
Electrostatic Precicpitators remove particulate matter as gas passes through a device that has charged metal plates. The particles are then removed because of static electricity. These systems can remove between 99-99.9% of particulate matter (Clean Coal Technologies).

====Super-critical Boilers====
Super-Critical Boilers and Ultra-Supercritical Boilers operate at temperature and pressure higher than regular boilers. By operating at higher temperature these systems become more efficient. Super Critical Boilers typically have 10%-20% CO2 emissions than other similar sub-critical technologies. Ultra-Supercritical boilers can be as much as 30% more efficent than sub-critical technologies (Clean Coal Technologies).

====More Information====

For more information regarding some of these separations equipment see [[Solids-involved equipment]]

For more information Regarding Cyclones see [[Separation processes#Cyclones| Cyclones]], and for modeling cyclones see their[[Solids-involved equipment#HYSYS Simulation|HYSYS Simulation]]

Methods for removing inorganic and organic gaseous pollutants include: absorption, adsorption, condensation, and combustion (Seider 2009 pg 609). A list of typical of gases pollutants and their sources from ''Plant Design and Economics for Chemical Engineers'' is shown below (Peters, 1991).

[[File:GaseousPollutants.png|thumb|border|center|middle|upright=4|link=|atl=| Common Gaseous Pollutants and their Sources]]

Typically Gas-liquid absorption processes are done completed in a vertical, countercurrent, flow through packed, plate, or spray towers. These systems require good liquid-gas contact and proper equipment. These systems also often have significant energy consumption because of large pressure drops (Peters, 2003). For high volume systems absorption by scrubbing with water or another solvent is the most widely used method (Towler, 2012). Dry adsorbents can be used to remove the last races of gaseous pollutants. Adsorption typically requires blowers, condensers, separators, and controls. You also typically need two packed beds so that one can be used while the other is regenerated. Examples of adsorbents are molecular sieves and activated carbon. Incineration is typically used when there are gas streams that have no recovery value. This can be done with direct flame or catalytic oxidation. Catalytic oxidation usually has higher capital costs, but lower operating costs because it does not require fuel.

You can find more information on [[Separation processes#Absorption| Absorption]] and [[Separation processes#Adsorption| Adsorption]] in [[Separation processes]]

===Outcomes===

====United States Common Pollutant Emissions====

The United States implemented the Clean Air Act in 1970 and since then emissions in the U.S. have been drastically reduced. As a matter of fact despite increases in the population in the last 35 years the amount of pollutants emitted have decreased by almost 70% (EPA). More importantly this demonstrates that reasonable efforts can be put towards environmental protection without causing too much harm to industry.

[[File:EPA.png|thumb|border|center|middle|upright=4|link=|atl=| GDP and Other Growth Factors Vs Common Pollutants Released]]

====China Clean Coal====

=====Success Through 2005=====

Coal is a very inexpensive and abundant source of energy and is abundant in China (Xu 2010). In China Coal the cause of 90% of SO2 emissions, 70% of dust emissions, and 67% of NOx emissions, and 70% of CO2 emissions. These numbers are staggering considering that Coal has never been more than 50 percent of China's Energy Supply (Xu 2010).

[[File:Coal.jpg|thumb|border|center|middle|upright=4|link=|atl=| Coal Consumption for Power and the Percentage of Coal Used (Xu 2010)]]

Despite the increasing coal consumption, high efficiency electric dust removal systems with efficiencies that are as high as 99.6% have greatly decreased soot emissions to 32% below 1980's levels as of 2005.

[[File:Soot.jpg|thumb|border|center|middle|upright=4|link=|atl=| SO2 and Soot Emissions from 1981 to 2005 (Xu 2010)]]

=====Technologies and Life Cycle Comparison=====
As of 2010 China was consuming 48.2% of coal globally. Four potential technologies integrated gasification combined cycle (IGCC), sub-critical coal power generation (Sub-C); super-critical coal power generation (Super-C) ultrasuper-critical coal power generation (USC). These technologies are compared mostly on the basis of net generating efficiency and efficiency. Net generating efficiency is the output of the plant divided by the total available energy in the fuel used (Liang, 2013).

where efficiency is define as: η=E/Etotal

and

Etotal=Emining+Etransportation+Egeneration

[[File:technologies.PNG|thumb|border|center|middle|upright=4|link=|atl=|A comparison of many of the parameters important to lfe-cycle analysis (Liang, 2013)]]

This demonstrates that USC has the best net generation efficiency, while also having the largest capacity for a single system.

[[File:Lifecycle2.PNG|thumb|border|center|middle|upright=4|link=|atl=| (Liang, 2013)]]

This breaks down the energy efficiency of each system and demonstrates that USC is the most efficient

[[File:Lifecycle.PNG|thumb|border|center|middle|upright=4|link=|atl=| (Liang, 2013)]]

This demonstrates that the capital cost per unit of energy production is very high for IGCC, but shows that Super-C and USC are very competitive.

[[File:Price.PNG|thumb|border|center|middle|upright=4|link=|atl=| Capital cost in $/kW of clean coal power-generation technologies (Liang, 2013)]]

Lastly IGCC and USC are lowest on different element, but since these all have different global warming potential it is difficult to tell which is the most efficient.

And the table below defines all the relevant terms.

[[File:GWP.PNG|thumb|border|center|middle|upright=4|link=|atl=| Environemntal measures and their explanations (Liang, 2013)]]

All of the data suggests that USC is the highest in both energy efficiency and net generating efficiency while competitive in price and the lowest on emissions among the different boilers.

==Solid Waste==
U.S. federal regulations require that solid waste be classified as hazardous or nonhazardous. Conditions for a classification of hazardous include: ignitability, corrosivity, reactivity, toxicity, or posing a substantial threat to the surrounding environment and its inhabitants. Hazardous waste must be treated on- or near-site before being removed in containers. Non-hazardous waste may be landfilled or incinerated in some cases. A typical estimate of costs for waste disposal is $0.03/lb for nonhazardous solids and $0.10/lb for hazardous solids. (Seider 2009 pg 609)

=References=

*Biegler LT, Grossmann IE, Westerberg AW. ''Systematic Methods of Chemical Process Design'' Prentice-Hall: Upper Saddle River, 1997.

*Broughton, J. ''Process Utility Systems: Introduction to Design, Operation and Maintenance'' Institution of Chemical Engineers: Rugby, Warwickshire, UK, 1994.

*Clean Coal Technologies. America's Coalition for Clean Coal Electricity. http://www.americaspower.org/clean-coal-technologies-1663/

*Duke Energy Company (2013). How Do Coal Fired Plants Work? Charlotte: Duke Energy.

*"Economizers." Energy Solutions Center Inc.: Boiler Consortium. http://www.cleanboiler.org/Eff_Improve/Efficiency/Economizers.asp. Accessed February 21, 2016.

*Lehman B, Worrell E. ''Electricity Production from Natural Gas Pressure Recovery Using Expansion Turbines.'' Lawrence Berkeley National Laboratory; 2001.

*"Overview of the Clean Air Act and Air Pollution." Environmental Protection Agency. November 17, 2015. http://www.epa.gov/clean-air-act-overview. Accessed February 5, 2016.

*"Oxygen Control." Energy Solutions Center Inc.: Boiler Consortium. http://www.cleanboiler.org/Eff_Improve/Efficiency/Oxygen_Control.asp. Accessed February 21, 2016.

*Peters, Max S.; Timmerhaus, Klaus D.; West, Ronald E. (2003). " Plant Design and Economics for Chemical Engineers." McGraw Hill Higher Education.

*Ruina Xu Energy (2010). "Clean coal technology development in China." Elsevier.

*Seider, Seader, Lewin, Widagdo. (2009). ''Plant Design and Economics for Chemical Engineers, 5th Edition.'' Hoboken: Wiley.

*Seider, Seader, Lewin. (2008). ''Product and Process Design Principles, 2nd Edition.'' Hoboken: Wiley.

*Summary of the Clean Water Act. United States EPA website. http://www.epa.gov/laws-regulations/summary-clean-water-act

*Towler, G.P. and Sinnot, R. (2012). ''Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design.'' Elsevier.

*Turton R.; Bailie, R.C.; Whiting, W.B.; Shaeiwitz J.A.; Bhattacharyya D. (2012). ''Analysis, Synthesis, and Design of Chemical Processes.'' Upper Saddle River: Prentice Hall.

*G.D. Ulrich, A Guide to Chemical Engineering Process Design and Economics, Wiley: New York, 1984.

*Xiaoye Liang, Zhihua Wang, Zhijun Zhou, , Zhenyu Huang, Junhu Zhou, and Kefa Cen. (2013). "Up-to-date life cycle assessment and comparison study of clean coal power generation technologies in China" Elsevier.

Separation processes

2016-02-22T01:44:01Z

Bts511: /* Cyclones */

Authors: Nick Pinkerton, [2014] Karen Schmidt, [2014] James Xamplas, [2014] Emm Fulk, [2015] and Erik Zuehlke, [2015] John Dombrowski [2016] , Brett Sleyster [2016] , Robert Cignoni [2016] , and Osman Jamil [2016] 

Stewards: David Chen, Jian Gong, and Fengqi You

Date Presented: February 9, 2014 /Date Revised: February 1, 2014

 

==Introduction==
Essentially all chemical processes require the presence of a separation stage. Most chemical plants comprise of a reactor surrounded by many separators. Separators have a countless number of jobs inside of a chemical plant. A separator can process raw materials prior to the reaction, remove incondensable gases, remove undesired side products, purify a product stream, recycle materials back into the process, and many other jobs that are essential to the process.

Chemical engineers must understand the science of separation and the variety of ways that separation can take place. There are many ways to perform a separation some of these including: distillation, absorption, stripping, and extraction. The science of separation revolves around the presence of two phases that are in contact and equilibrium (Wankat, 2012).

[[File:Sepmeth.JPG|frame|Figure 1. Separation methods by property]]

==Theory==
===Vapor-Liquid Equilibrium===
Separation processes are based on the theory of vapor-liquid equilibrium. This theory states that streams leaving a stage in a separation process are in equilibrium with one another. The idea of equilibrium revolves around the idea that when there is vapor and liquid in contact with one another they are in constantly vaporizing and condensing. Different components in the mixture will condense and vaporize at different rates. There are three types of equilibrium conditions that can be subdivided into thermal, mechanical and chemical potential categories. These separate equilibrium states are given as:

<math>T_{liquid} = T_{vapor}</math>

<math>p_{liquid} = p_{vapor}</math>

<math>chemical potential_{liquid} = chemical potential_{vapor}</math>

==Distillation==
===Flash Distillation===
Flash Distillation is one of the simpler separation processes to be employed in a chemical plant. The main premise of flash distillation is that a portion of a liquid feed stream vaporizes in a flash chamber or a vapor feed condenses. Vapor-liquid equilibrium will cause the vapor phase and the liquid phase to have different compositions. The more volatile component of the mixture will compose of a larger portion of the vapor. This simple separation is easy to manufacture but does not result in large degrees of separation.

Flash distillation requires a feed stream that is pressurized and heated and then passed through a valve into a flash drum. The large pressure drop across the valve will result in a partial vaporization of the fluid. Vapor will be removed overhead from the flash drum while the remaining liquid will collect at the bottom of the drum and be removed. Most flash drums will contain an entrainment eliminator which is a screen that prevents liquid from being carried into the vapor effluent. Figure 2 shows a simple overview of the flash distillation process. As shown, there is a heater that flows into a let-down valve where the two-phase flow begins. Variables y and x are the mole fractions of the more volatile component in the vapor and liquid effluents, respectively.

[[File:Flash.gif|center|frame|Figure 2. Flash Distillation Flow Diagram]]

===Column Distillation===
Distillation columns are the most widely used separation technique used in the chemical industry, accounting for approximately 90% of all separations (Wankat, 2012). Distillations in columns consist of multiple trays that each act at their own equilibrium conditions. Large columns are able to perform complete separations of binary mixtures as well as more complex multi-component mixtures.

[[File:column.jpg|250px|center|]]
===Stages===
Columns are separated into stages by the presence of trays. These trays allow for vapor-liquid contact and equilibrium to occur. Typically, the more stages in a column, the larger separation that can be achieved. There are many different types of trays that can be used in a column.
====Sieve Trays====
The simplest and least expensive tray type is the sieve tray which is a sheet of metal with holes punched into it to allow vapor flow. Sieve trays can have different hole patterns and sizes that will affect the tray efficiency and flow rates.

[[File:sieve.jpg|200px|center|]]

====Sieve Tray Design Procedure====

The design of these plates is done through a trial-and-error process. Most commercial process simulations (such as HYSYS) have default tray designs, and automatically specify dimensions. However, these dimensions selected or calculated by the simulations may not give the best performance for your system, so it is valuable to understand how to design the sieve trays and how specific parameters may affect performance. Hand calculations using the following methods can be used to guide the simulation programs to better design. This section will use sample data to work through an example of the process. The following is a general list of steps for designing a sieve plate:

=====1. Calculate the maximum and minimum vapor and liquid flow rates for the turndown ratio required.=====
This data can be collected from a McCabe-Thiele diagram and/or from process simulation data.

Data from McCabe Thiele diagram, for example: 
Number of stages = 10 
Slope of top operating line = 0.185 
Slope of bottom operating line = 1.43 
Top composition = 98.8 mol% acetone 
Bottom composition = 4 mol% acetone (a) 
Minimum reflux ratio = 0.31 

Collect data rom mass balances: 
MW of feed = (0.6 kmol a/kmol)*(58 kg/kmol a) + (0.4 kmol acetic acid/kmol)*(60 kg/kmol aa) = 59 kg/kmol 
Feed rate = 100 kmol/h 
Using material balances... D = 57.3 kmol/h ; B = 42.7 kmol/h 
Vapor rate, V = D(1+R) = 57.3 kmol/h (1+.31) = 75.1 kmol/h 
Below feed line, slope of BOL = 1.43 = L'm/V'm (liquid rate/vapor rate)

=====2. Collect or estimate the system physical properties.=====
Here it is important to know information about both the top and bottom of the column. Useful information includes temperature, pressure, column pressure drop (a common assumption is 100 mmH2O per plate), densities, molecular weights, surface tensions, and number of stages (which can be estimated from the McCabe-Thiele diagram).

=====3. Select a Trial Plate Spacing=====
The plate spacing will depend on the column diameter and operating conditions. Plate spacings from 0.15 m to 1.0 m are typically used. The smaller the diameter, the smaller the spacing. Small columns will use close spacing. Columns with diameters above 1.0 m, plate spacings of 0.3 m to 0.6 m are normally used. A good initial estimate is 0.5 m.
[[File:trayspacing.jpg|400px|center|]]

=====4. Estimate the column diameter, based on flooding considerations.=====
Vapor and liquid flow rates will vary along the column, so plate design needs to be considered both above and below the feed. Using plate spacing and FLV (which is the square root of the ratio of the liquid to vapor flow rates), you can obtain the value of K from the plot.

[[File:floodingplot.jpg|400px|center|frame|Figure. Plate Spacing]]
There is a range of vapor and liquid flow rates in which the column needs to be operated. Too low or too high of rates can result in various inefficiencies in the column operation, as shown in the figure below. For example, if the vapor rate is too high, flooding will occur. However, it is not safe to operate on the flooding line. Instead, columns are typically designed for 80% of flooding at the maximum flow rate.
[[File:vap_rate_vs_liq_rate.jpg|400px|center|frame|Figure. Tray behavior]]

=====5. Decide the liquid flow arrangement.=====
Common flow arrangements are single pass (cross flow), double pass, and reverse flow. Using conditions at the bottom of the column, calculate the max volumetric flow rate. Use this flow rate and the column diameter to determine the preferred flow arrangement from the chart below.
[[File:Liquidflow.jpg|200px|center|]]

=====6. Make a trial plate layout: downcomer area, active area, hole area, hole size, weir height.=====
Standard sizes for trays -- and good assumptions for the first iteration -- are: weir height, hw = 50mm ; hole diameter, Dh = 5mm ; plate thickness, tpl = 5mm. From the graph below, the ratio of downcomer area (Ad) to column cross-sectional area (Ac) can be determined from the ratio of weir length (lw) to column diameter (Dc) and vice versa.
[[File:platelayout.jpg|200px|center|frame|Figure. Plate Dimensions]]

=====7. Check the weeping rate=====
Compare the actual vapor velocity to the minimum vapor velocity -- if velocity is too low fluid will "weep" through the tray holes. If the weeping rate is unsatisfactory, return to step 6 and choose different values for the plate layout dimensions. From the chart in step 4, it can be seen that there is a minimum vapor flow rate below which the liquid "weeps" from the tray above.

For the remaining steps in this design process, it is recommended to check your assumptions after each step and revise them as necessary in order to maintain operation in the "sweet spot" of the vapor rate vs. liquid rate plot. Additional iterations may be required as you move through the procedure.

Calculate the maximum liquid flow rate. Calculate the minimum liquid flow rate at 70% turndown (recommended). Calculate the height over the weir as
<math>h_o=750[\frac{L_w}{p_Ll_w}]^\frac{2}{3}</math>

=====8. Check the plate pressure drop=====
<dfn>If the pressure drop calculated here is too high, return to step 6.</dfn>

Proceed to step 9 if the pressure drop assumption is valid.
=====9. Check the downcomer backup. =====
<dfn>If the downcomer backup is too high, return to step 6 or 3.</dfn>

The plate spacing affects the amount of fluid in the downcomer. Calculate the level in the downcomer and the residence time of the fluid to see if the values are valid. Note that residence times greater than 3 seconds are acceptable.

Proceed to step 10 if residence time is acceptable.
=====10. Decide plate layout details.=====
Determine calming zones, the unperforated areas at the inlet and outlet sides of the plate. The width of each zone is usually made the same. Recommended values are: below 1.5 m diameter, 75 mm; above, 100 mm. The unperforated area can be calculated from plate geometry. Also check the hole pitch, or the distance between hole centers. It should not be less than 2.0 hole diameters. A normal range is between 2.5 and 4.0 hole diameters. The shape must also be specified. Square and equilateral triangle holes are used.

=====11. Recalculate the percentage flooding based on the chosen column diameter.=====
An assumption of 80% flooding was chosen so that operation would occur in the "sweet spot." This assumption must be checked by calculating the flooding percentage for a given column diameter.
uv = (max volumetric flow rate)/(net area)

<math>%flooding = \frac{u_v}{u_f}</math>

If the hole pitch is unsatisfactory, return to step 6.
=====12. Check entrainment=====
''If too high, return to step 4'' Use the graph below to determine entrainment from FLV.
[[File:entrainment.jpg|400px|center|]]

The value for fractional entrainment can be used to re-estimate the column efficiency, and reevaluate the number of trays needed. Can return to step 1 for more accurate estimates.

=====13. Optimize design.=====
After returning to step 1 to reevaluate the number of trays, it is valuable to repeat steps 2 through 12 to find the smallest diameter and plate spacing acceptable at the lowest cost.

=====14. Finalize the design.=====
Optional: draw up the plate specification and sketch the layout of the plate.

====Bubble-Cap Trays====
Bubble-cap trays consist of a weir around each hole in the tray which is covered with a cap that has holes or slots to allow vapor passage. Entrainment is about three times larger than a sieve tray. Bubble-cap trays require larger tray spacing than sieve tray design. Bubble-cap trays have been known to have problems with coking, polymer formation, or high fouling mixtures. Recently, very few new bubble-cap columns are being built due to the expense and marginal benefits. However, engineers will likely encounter bubble-cap columns still currently in operation.

====Flow Patterns====
Cross flow columns are the most common pattern for distillation columns. For liquid flows between 50 and 500 Gal/min, a cross flow column is appropriate. When liquid flow is increased above 500 Gal/min, an engineer should consider designing a double pass or multi-pass column. This will reduce the liquid gradient on the tray and reduce the downcomer loading (Wankat, 2012).

===Column Sizing===
Column height will be dependent on the amount of trays required and the spacing between the trays. Normally, tray spacing of 0.15 m to 1 m is used. For columns, above 1 meter in diameter, 0.5 m can be used as an initial estimate.

Column diameter is influenced by the vapor flow rate in the column. The trays can not have excess liquid entrainment or high pressure drops; therefore, vapor velocity in the column must be maintained at a reasonable level.

An equation based on the Souders and Brown equation can be used as an estimate for the max allowable superficial vapor velocity,

<math>\hat u_v = (-0.171l_t^2 + 0.27l_t - 0.047){\frac{\rho_L - \rho_v}{\rho_v}}^{1/2}</math>

where <math>l_t</math> is the plate spacing in meters, <math>\rho_L</math> is the density of the liquid stream, and <math>\rho_V</math> is the density of the vapor stream.

Column diameter, <math>D_c</math>, can then be estimated using the relation,

<math>D_c = \sqrt{\frac{4\hat{V_w}}{\pi\rho_v\hat{u_v}}}</math>

where <math>\hat{V_w}</math> is the maximum vapor rate in kg/s (Towler et al., 2013).

===Distillation Applications===

Distillation is a process that can be implemented in various scales. There is both laboratory scaled distillation as well as very large industrial distillation. Other applications for distillation include food/alcohol processing and herb distillation for the perfume and medical industries. Typically laboratory scaled distillation occurs in batches whereas industrial distillation (e.g. fractional distillation of crude oil) occurs continuous with a constant distillate and bottom effluent streams.

Some applications of distillation are concerned the top stream only, some the bottom stream only and others both streams can be used for future products. In alcohol distillation for example, the water that is separated from the ethanol/water binary solution is discarded as waste water. In fractional distillation of crude oils, the heavy hydrocarbons at the bottom of the column are collected and sold along with the light hydrocarbons that appear in higher side draws (Wankat, 2012).

===Example Case: Ideal Distillation===

Assume an equimolar mixture flowing at 10 mol/s of 20 mol% n-pentane, 30 mol% n-hexane, and 50 mol% n-heptane. Separate the mixture into 3 products: 99% pure n-pentane, 99% pure n-hexane, 99% n-heptane. Assume the feed and products are all liquids at the bubble points. There are two process alternatives to consider in this example. The direct sequence removes the most volatile species, pentane, in the first column, and then separates hexane and heptane in the second column. The indirect sequence separates the heaviest product, heptane, and then separates pentane from hexane in the second column. This example will consider the direct sequence. Next, we must decide if these species exhibit fairly ideal behavior during distillation. Since the n-alkanes have very similar properties, it is safe to assume they will display close to ideal behavior. The next step is to look up the boiling points of the 3 species. In this case, the normal boiling points of pentane, hexane, and heptane are 309 K, 342 K, and 372 K, respectively. Also, it is a good idea to look up relative volatilites, to further verify near-ideality of the mixture, but also to obtain the information necessary for the Underwood method, which we will employ to obtain a solution. The next step is to write out material balances based on molar flows and the design specifications. They go as follows:

<math>\mu_I(nC5) + \mu_{II}(nC5) = 2 mol/s</math>

<math>\mu_I(nC6) + \mu_{II}(nC6) + \mu_{III}(nC6) = 3 mol/s</math>

<math>\mu_{II}(nC7) + \mu_{III}(nC7) = 5 mol/s</math>

<math>\mu_I(nC5) = 99\mu_I(nC6)</math>

<math>\mu_{II}(nC5) = (5/990)\mu_{II}(nC6)</math>

<math>\mu_{II}(nC7) = (5/990)\mu_{II}(nC6)</math>

<math>\mu_{III}(nC7) = 99\mu_{III}(nC7)</math>

where <math>\mu</math> represents the molar flow, and the subscript represents the product stream.

Solving this system of equations:

<math>\mu_I(nC5) = 1.985\ mol/s</math>

<math>\mu_{II}(nC5) = 0.015\ mol/s</math>

<math>\mu_I(nC6) = 0.020\ mol/s</math>

<math>\mu_{II}(nC6) = 2.930\ mol/s</math>

<math>\mu_{III}(nC6) = 0.050\ mol/s</math>

<math>\mu_{II}(nC7) = 0.015\ mol/s</math>

<math>\mu_{III}(nC7) = 4.985\ mol/s</math>

At this point we have enough information to use Underwood's method to estimate the minimum vapor flows in the column. The following three equations are used in Underwood's method:

<math>\sum_i \frac{\alpha_{ik}}{\alpha_{ik}-\phi}f_i = (1-q)F</math>

<math>(R_{min}+1)D = \sum_i \frac{\alpha_{ik}}{\alpha_{ik}-\phi}d_i = V_{min}</math>

<math>\bar R_{min}B = -\sum_i \frac{\alpha_{ik}}{\alpha_{ik}-\phi}b_i = \bar V_{min}</math>

where <math>\alpha_{ik}</math> is the relative volatility of species <math>i</math> to species <math>k</math>, <math>f_i</math> the molar flow of species <math>i</math> in the feed, <math>q</math> the fraction of the feed that joins the liquid stream at the feed tray, <math>F</math> the total molar flow of the feed, <math>D</math> the molar flow of the distillate, <math>R_{min}</math> the minimum reflux ratio <math>(=L_{min}/D)</math>, <math>d_i</math> the molar flow of species <math>i</math> in the distillate, <math>V_{min}</math> the minimum vapor flow possible in the top section of the column to accomplish the desired separation, <math>\bar R_{min}</math> the minimum reboil ratio <math>(=\bar V_{min}/B)</math>, <math>b_i</math> the molar flow of species <math>i</math> in the bottoms product, and <math>\bar V_{min}</math> the minimum vapor flow in the bottom section of the column. The final variable, <math>\phi</math>, will be solved for using the first Underwood equation, and it's value will be decided based on the relative volatilities of the key components in the column.

So, after solving the first Underwood equation, we get two values for <math>\phi</math>, 3.806 and 1.462. Because 3.806 is between the relative volatilities of the key components, we will substitute that value for <math>\phi</math> into the second Underwood equation. Doing so for both columns gives <math>V_{min} = 6.4\ mol/s</math> for the first column and <math>V_{min} = 8.9\ mol/s</math> for the second column, for a total minimum vapor flow of 15.3 mol/s. The process would then be repeated for the indirect sequence, and the decision for which process to use would be justified by the process with the overall minimum vapor flow (Biegler et al., 1997).

==Absorption==
===Description of Absorption===
Another separation process used in industry is absorption, which is used to remove a solute from a gas stream. It accomplishes this by contacting the gas mixture with a liquid solvent that readily absorbs the undesirable components from the gas stream, purifying the gas stream. This separation process is determined by the inputs of the liquid flow rate, temperature, and pressure.

The absorption factor, which can be determined mathematically, determines how readily a component will absorb in the liquid phase. The absorption factor of component i is

<math>A_i=L/K_iV</math>

where <math>L</math> is the liquid flow rate entering the column, <math>V</math> is the vapor flow rate entering the column, and <math>K_i</math> is the vapor/liquid equilibrium ratio for component i (Peters & Timmerhaus, 2003). Higher absorption factors result in higher absorptivity into the liquid and a decrease in the number of trays required for separation, however a diminishing return occurs after the absorption factor is greater than 2.0. An absorption factor of 1.4 is most commonly used.

In general absorption can be seperated into two overarching categories, physical and chemical absorption. In physical absorption, the unwanted solute in the gas is absorbed into the liquid phase because solubility of the component is higher in the liquid phase than the gas phase. In chemical absorption the solute is removed from the gas via a reaction with the solvent, this reacted product is then transported into the liquid phase (Danckwerts 1965). There are two types of chemical absorption reversible and irreversible. Generally reversible chemical absorption is preferred as the solvent can be put through a stripper and regenerated so it can be recycled back to the absorption process (Wankat, 2012).

===Absorption Apparatus===

There are five major apparatus used for absorption in industrial application. These five pieces of equipment are spray absorbers (or towers), ejector (venturi) scrubbers, packed columns, trayed columns, and film absorbers (Schmidt, 2012).

==== Spray Tower vs Ejector Scrubber ====

In both '''spray tower''' and the '''ejector scrubber''' nozzles are employed to produce small solvent droplets. These small droplets increase the surface area of the liquid to gas contact allowing for the maximum amount of mass transfer to occur between the gas mixture and the liquid. The major difference between the two nozzle equipment designs is the configuration and type of nozzles. In the ejector scrubber shown in Figure 3 there is a single nozzle that is generally a higher pressure spray nozzle that produces finer solvent drops allowing for an even greater amount of mass transfer enabling better physical absorption (Schmidt, 2012).
[[File:Ejectorventuri.jpg|thumb|200px|center|Figure 3. Ejector Scrubber (US EPA, 2006)]]
'''Spray towers''' on the other hand generally have many nozzle at different heights where the liquid solvent will be sprayed out of to contact the gas running through the tower. This design is used in order to ensure the gas contacts the liquid as throughout the tower. These nozzles are lower pressure than a ejector scrubbers nozzle and thus physical mixing is worse in this configuration. Since physical mixing is generally worse in this configuration it is usually used in conjunction with a chemical absorption process. The other major difference between the ejector scrubber and the spray tower is that gas and liquid flow is cocurrent in the former while it is countercurrent in a spray tower. A spray tower absorber is shown below in Figure 4 (Schmidt, 2012).
[[File:SparyTowerAbsorber.jpg|thumb|200px|center|Figure 4. Spray Tower Absorber (US EPA, 2006)]]

==== Tower Type Absorption Apparatus ====
'''Packed column absorbers''' and '''tray column absorbers''' have very high efficiencies for the removal of an unwanted solute in the gas stream. The major disadvantage a trayed column has when compared to a packed column is the pressure drop. The pressure drop in a packed column is generally very low, whereas in between each tray of a trayed column pressure drop can be quite large. However the advantages inherent to trayed columns become clear when one needs the solvent to have a high concentration of the component to be removed from the gas stream. This is most important in the case where there is a very low concentration of the component in the gas stream and the specification states the solvent must contain a high concentration of that component. In this case the flow rate of the solvent may not be high enough for a packed column, however in a trayed column the solvent flow rate can be near zero for operation (Schmidt, 2012). Packed and trayed column internals are very similar to the setups found in the respective distillation columns.

For a '''trayed column''' the plate efficiency can be calculated using O'Connell's Correlation which invovles the Henry's Law constant, total system pressure, and solvent viscosity at the operating temperature (Towler & Sinnott, 2013).

<math>x=0.062*\frac{\rho_s*P}{\mu_s*H*M_s}</math>

where
<math>x</math> is the tray efficiency,
<math>\rho_s</math> is the density of the solvent in <math>kg/m^3</math>,
<math>P</math> is the total pressure of the system in <math>N/m^2</math>,
<math>\mu_s</math> is the solvent's viscosity in <math>mNs/m^2</math>,
<math>H</math> is the Henry Law constant in <math>1/(Nm^2*(mol fraction))</math>,
and <math>M_s</math> is the molecular weight of the solvent.

A packed towers height can be determined using the equations below when concentration of solute is below 10% so that the assumption that the flow of gas and liquid will be essentially constant throughout the column holds (Towler & Sinnott, 2013). The height of packing <math>Z</math> is given by the following equation:

<math>Z=\frac{L_m}{K_G*a*P}*\int\limits_{y_2}^{y_1} \frac{dy}{y-y_e}\,</math>

where <math>P</math> is the total pressure, <math>a</math> is the interfacial surface area per unit volume, <math>y_1</math> and <math>y_2</math> are the mol fractions of the solute in the gas stream at the bottom and top of the column respectively, <math>G_m</math> is the molar gas flow rate per unit cross-sectional area, and <math>y_e</math> is the mole fraction of solute in the gas that would be in equilibrium with the liquid concentration.

The first half of the equation before the integral can be called the height of an overall gas-phase transfer unit <math>H_G</math> and the second part of the equation is the number of overall gas-phase transfer units or <math>N_G</math>. Using these definitions the above equation can be simplified to

<math>Z=H_G*N_G</math>

These equations assist in sizing an absorption column (Towler & Sinnott, 2013).

==== Film Absorber ====
The final absorber the film absorber is generally used in the case where the heat of absorption must be removed. The film absorber operates by sending the gas and solvent through a heat exchanger where the solvent creates a thin film on the walls of the tubes and the gas flows through the interior allowing for solute transfer. The good heat transfer present in a film absorber makes it preferable for situations where low temperatures are required for a high recovery of the solute (Schmidt 2012).

===Industrial Absorption Processes===
An industrial example is lean oil absorption, which is used to separate nitrogen and other impurities from natural gas. A lean oil is contacted with low quality natural gas, and the methane is selectively absorbed by the lean oil, leaving the impurities behind. The methane is subsequently regenerated from the rich oil as high quality natural gas (Petrogas Systems, 2014).

Other common industrial practices of absorption come from sour gas treatment. Amine gas treating is used to remove hydrogen sulfide or carbon dioxide from gas streams via a reversible chemical absorption. In amine gas treating the sour gas is fed to the bottom an absorber where amine solution is fed to the top along with any necessary make up water. The sour gas components are absorbed into the amine via a chemical absorption method. Sweet gas leaves the top of the absorber whereas the amine out of the bottom, now rich with acidic components is sent to a regenerator where the acid gas components are stripped and the acid gas is generally sent to a flare whereas the amine now lean again is recycled back into the first absorber (Miller & Zawacki, 1978). Figure 5 below shows the typical setup of an amine plant. Another type of sour gas treatment that uses absorption is Merichems LO-CAT process which uses a chelated iron to remove hydrogen sulfide from feed gas in the absorption column (Merichem 2015).
[[File:AmineTreating.png|thumb|400px|center|Figure 5. Amine Gas Treating Plant Schematic]]

==Stripping==
This process separates solutes from solvents (often after absorption, to purify the solvent so that it can be recycled to an absorber). Stripping will depend on the vapor and liquid flow rates, as well as the temperature and pressure of the column. There is a temperature drop down the column, so columns generally have either an increased operating temperature or decreased operating pressure.

The stripping factor of component i is

<math>S_i=K_iV/L</math>

where <math>K_i</math> is the vapor/liquid equilibrium ratio, <math>V</math> is the vapor flow rate entering the column, and <math>L</math> is the liquid flow rate entering the column, will determine how much of solute i will be stripped from the liquid into the vapor phase (Peters & Timmerhaus, 2003). The usual range for the stripping factor is between 1.2 and 2.0, with a stripping factor of 1.4 being most economic.

An example of stripping in industry is the deodorization of food items such as oils. The oil is heated and allowed to trickle down the column while steam flows up from the bottom of the column. At the vapor-liquid interface, volatile components of the oil transfer to the steam and are carried off the top of the column, leaving a purified oil product (Alfa Laval, 2014).

==Bioseparations==
===Importance===
As our ability to manipulate and engineer biological systems improves, biological products are becoming an increasingly important source of therapeutics and fuels. The production of fuels from biomass via either the enzymatic breakdown of a feedstock or the secretion of usable lipids from algae is a promising new energy source. Additionally, enzymes, antibodies and other therapeutic proteins have been applied to the treatment of a wide range of diseases. Although each process requires its own set of separations, all follow the same basic format: separation of biomass, product isolation, and product purification (Belter et al., 1998). This section will provide examples of unit operations in each step. Ultimately, the choice of separation process and unit operations will depend on the specific process and product. The descriptions below are examples of the most common bioseparation operations within the general platform (Harrison et al., 2003).

Bioprocesses begin with fermentations or growth operations. In biofuel production processes, this may involve growing algae or breaking down corn or cellulosic biomass. For the production of therapeutics, mammalian or bacterial cells may be grown in a fermentor and the product secreted into the supernatant or harvested from the cells.

===Biomass Separations===
After fermentation and product production, the solid biomass must first be separated from the desired product. If the product is secreted from the cells, this can be done immediately after fermentation ends. If the product is not secreted, the cells must first be lysed.
Cell lysis is the process of lysing, or breaking, the cell in open. Mechanical lysis is the simplest, and involves physically breaking the cell either by mashing (think mortar and pestle) or blending the cells into a homogenous solution in a homogenizer. Chemical lysis is another method, achieved by introducing an osmotic shock or chemically degrading the cell membrane. Additional separation can be achieved by flocculation, which is the process of aggregating biomaterial by charge neutralization or bridging. These larger complexes are easier to separate from smaller molecules (Harrison et al., 2003).

The next step is removing the unwanted biomass from the product in solution. Separation by centrifugation or sedimentation are the most common, although filtration is sometimes also used for processes where a biomass cake is desired. Both methods utilize density differences to separate the product from the solid biomass (Towler and Sinnott, 2013).

====Sedimentation====

Sedimentation relies purely on the force of gravity, while centrifugation speeds the settling process by subjecting the cells to a centrifugal force. Sedimentation in a settling tank is the simplest method of solid-liquid bioseparation. In this process, biomass in a tank is simply allowed to settle to the bottom over time. While this process is inexpensive, requires little energy and can separate out large volumes of biomass, it generally requires long time periods and is only mostly in very large-scale processes where active centrifugation is difficult (Belter et al., 1998).

====Centrifugation====
Centrifuges are widely utilized across many processes, and thus a wide variety of scales and designs have been developed. Disk-stack centrifuges, in which the solid phase is deposited onto “shelves” in the center of the spinner and liquid phase is pushed to the outside, are some of the most commonly used centrifuges in industry. They are especially suited to biomass separation processes because they can be built on a large scale and are ideal for separating fine solids from liquids. [[File: Disk_stack_centrifuge_towler.png|frame|center|Fig. 6: Diagram of a disk-stack centrifuge (Tolwer et al, 1997).]] Tubular bowl centrifuges are also common and can reach separation efficiencies of up to 90%. Heavier products accumulate along the sides of the bowl, while the light phase flows out the top. They separate products by can be used both to separate solids from liquids and immiscible liquids, such as and oil product and an aqueous broth (Tolwer and Sinnott, 2013). [[File: tubular bowl centrifuge towler.png|frame|center|Fig. 7: Diagram of a tubular bowl centrifuge centrifuge (Tolwer and Sinnott, 2013).]]

Centrifugation scale-up is made easier by sigma analysis, which allows for the estimation of appropriate feed rates for different size centrifuges. The sigma factor is dependent on the inner and outer radius of the centrifuge, the angular velocity, and the sedimentation velocity of the solid particles being separated. It can be thought of as the characteristic cross-sectional area with units of [length]2. The sedimentation velocity can be calculated by

<math>v_g={\frac{2a^2(\rho-\rho_0)}{9\mu}}g</math>

where <math>v_g</math> is the sedimentation velocity, <math>a</math> is the cell or biomass particle diameter, <math>\rho</math> is the particle density, <math>\rho_0</math> is the fluid density, and <math>\mu</math> is the fluid viscosity. The volumetric flow <math>Q</math> can be estimated by

<math>Q=(v_g)(\Sigma)</math>.

The equality

<math>{\frac{\Sigma_1}{\Sigma_2}}={\frac{Q_1}{Q_2}}</math>

can be an easy way to estimate equivalent flow rates between a small-scale centrifuge 1 and larger centrifuge 2 (Harrison et al., 2003).

====Example: Centrifugation Scale-up====

You are trying to separate a cell of radius 0.4 <math>\mu</math>m with a density of 1.05 g/cm3 from broth of mostly water (density of 1 g/cm3 and viscosity of 0.01 g/cm s). The sigma factor of the centrifuge you are using is 1 x 106 cm2. A] What volumetric flow rate should you use? B] If you want to scale up the process to a centrifuge with <math>\Sigma</math> = 3 x 106 cm2, what flow rate would you use in the larger centrifuge?

Solution:
A] Using the equation for <math>v_g</math>, and being mindful of units, the sedimentation velocity equals 1.74 x 106 cm/s. The flow rate, then, equals

<math>Q=(1.74 x 10^-6)(1,000,000) = 1.74 cm^3/s = 0.104 L/min</math>.

B] Keeping in mind that for the same process, <math>v_g1 = v_g2,</math> and rearranging the sigma factor equality, the new flow rate is

<math>Q_2 = {\frac{\Sigma_2 x Q_1}{\Sigma_1}} = {\frac{(3 x 10^6)(0.104)}{1 x 10^6}} = 0.313 L/min </math>

===Product Isolation===
Liquid-liquid separation, to extract the product from the aqueous phase, is much less straightforward than liquid-solid extraction. Many methods - especially adsorption, filtration, and precipitation - are similar in principle to operations found in other, non-biological separations. The exact separations used depend on the nature of the product and the scale of the process. These processes are nearly identical to their non-biological counterparts, and their description is left to other sections.

Particular care needs to be taken with protein products because of their instability, and the selection of an appropriate solvent or adsorbent is crucial to a successful process (Harrison et al., 2003).

===Product Purification===
The final steps of protein purification and polishing remove any remaining contaminants and bring the concentration of product to an appropriate value for applications. Purification processes for food-grade and medical products can be extensive, as sterility and high purity are essential. Purification in fuel-producing processes may be less extensive, depending on the process. Chromatography and crystallization are two common steps in purification and are especially used in industrial scale protein production. Several different types of chromatography exist with the ability to carry out different types of separations.
Chromatography is similar to adsorption in that it relies on differences in affinity between solutes and a solid surface. A solution is eluted through a column containing a solid resin with various affinities for the substances in solution. In adsorption, the solutes are evenly saturated throughout the column. Chromatography differs in that solutes are deposited a resin phase before the column is flushed with an elution solvent specific that results in solutes eluted in bands.

==== Ion Exchange Chromatography ====

There are two main types of ion exchange columns—anion and cation. Anion exchange resins have a positive charge and are used to retain products with a negative charge. Cation exchange resins have a negative charge and are used to retain products with a positive charge. The pH of the elution buffer is change to force a specific solute to wash out, depending on whether the pH of the buffer is above or below the isoelectric point of the solute (Belter et al., 1998). This is especially useful for the separation of protein product (including antibodies), nucleic acids, and other charged molecules. When the solutes have sufficiently different isoelectric points, the pH of the buffer is manipulated to affect the solute charge and force the product to elute while the solute remains preferentially bound to the resin, or vice versa (Harrison et al., 2003). In general, the most strongly charged molecules will remain in the column for a longer period of time. Elution washes through the weakly bound ions before the more strongly bound ions. Different speeds of elution can be visualized as in figure 8.

[[File:chromatography.png|frame|center|Fig. 8: Illustration of product bands in an elution chromatography column (Belter et al., 1998).]]

==== Size Exclusion Chromatography ====

In gel filtration chromatography, small molecules are "trapped' by the porous resin and take longer to flow through the column. Larger products will elute first because the smaller molecules are better able to penetrate the resin. This forces them to take a much longer path through the column, which means it takes longer for them to elute. This operation is often used when there is a distinct difference in size between the desired product and other solutes.

==== Affinity Separations ====

Affinity chromatography is very similar to ion exchange chromatography in that the interactions between the material in the column and the molecules in the feed. The main difference is that affinity chromatography can rely on a great variety of types of interactions. Two very common types of affinity are exploited in affinity chromatography columns. The first is immunoaffinity. Proteins are specifically bound by antibodies which can be incorporated onto beads and used in chromatography. Antibodies are designed to bind only a single protein, so these interactions are considered to be highly specific. The protein can be eluted using a buffer that changes the pH or salinity in the column, which adversely affects binding (Hage, 1999).

Proteins can be separated from very complex and unrefined mixtures using this method because of the great specificity. Antibody-protein interactions are so specific that individual proteins can be isolated from blood samples as shown in the figure below. The main drawback of this method is simply that specific interactions between proteins and binding targets do not always exist. Antibodies exist for many proteins, but for others they must be customized or created using some sort of evolution process. This is not a trivial task and can require a significant capital investment (Hage, 1999).

[[File:Affinity_Chromatography_Example.PNG|frame|center|Fig. 9: Affinity purification separating fibrinogen from human plasma using an anti-fibrinogen antibody (Hage, 1999).]]

The other main type of affinity chromatography is based on protein specific tags and the molecules or surfaces to which they bind. One of the most common types of protein tags used is the polyhistidine tag. This tag consists of 6-8 consecutive histidine residues which can be added to the exterior of the desired protein product. The addition of this tag requires alterations to the coding sequence of the protein. The polyhistidine tag binds strongly with nickel and cobalt ions. The product with the tag can then be eluted with imidazole—a small molecule with the same structure as the functional group of the amino acid histidine. Imidazole will bind the cobalt and nickel ions more strongly than the histidine in the tag. Along with chromatography, protein tag interactions can be leveraged with the use of beads that can be deposited directly into the solution containing the protein of interest (Lichty et al., 2005).

Different proteins can be purified in this manner with varying levels of efficiency. There is a very high dependence on the size of the protein, electronic properties, and steric considerations. The figure below demonstrates varying separation efficiencies with proteins of different sizes (HisPur Cobalt Resin - Thermo Fisher Scientific, 2016).

[[File:HisTagPurification.PNG|frame|center|Fig. 10: Protein gel demonstrating the separation of proteins from a sample. The far left column represents the most efficient separating conditions using cobalt and the far right column represents a negative control with no purification performed (HisPur Cobalt Resin - Thermo Fisher Scientific, 2016).]]

Several other types of tag-bead interactions can be utilized in separations processes. Maltose Binding Protein is a small protein that can be added to a protein of interest. It binds strongly with beads coated in immobilized maltose and can be released by flushing with maltose. As MBP is a full sized protein that typically must be removed from the protein of interest in order for it to be used. In this case, the site specific TEV protease is often used cleave MBP from the protein of interest. In addition, under specific circumstances, other unique tags can be used and provide varying levels of specificity in separations. The Flag tag, 3x Flag tag, Glut tag, and Strep tag. While these are all commonly used, the polyhistidine tag is the most popular because it gives the highest level of specificity. Some of the main purification tags are compared in the table below (Lichty et al., 2005).

{| class="wikitable"
|-
! Tag
! Size (aa)
! Resin
! Eluting Agent
! Capacity (mg/ml)
! Cost/10 mg
|-
| His
| 6
| Talon
| Imidazole
| 5-14
| $18
|-
| Maltose Binding Protein
| 396
| Amylose
| Maltose
| 3
| $12
|-
| Glutathione S-Transferase
| 218
| GSH-Sepharose
| Glutathione
| 10
| $36
|-
| Strep II
| 8
| Strep-Tactin Sepharose
| Desthiobiotin
| 0.1
| $293
|-
| FLAG
| 8
| Anti-FLAG M2 MAb Agarose
| Flag peptide
| 0.6
| $1045
|}

Unfortunately, based on the contents of the solution, the characteristics of the protein, and the specific product being used, there are extremely large amounts of variation in yield and purity for each type of tag. Therefore, it is crucial to allocate time towards preliminary tests and scale-up experiments with the exact solution from which the protein of interest is being separated. The his tag is typically the most used because, as seen in the table above, it has one of the lowest costs and highest yields. In addition, because it is only a 6 amino acids, it often is not cleaved from the protein of interest which minimizes the number of steps required in the overall purification process (Lichty et al., 2005).

==== Crystallization ====

Crystallization, or the formation of solute crystals from a solution, is especially useful in biomolecule separations because it is possible to obtain a 99.9%+ product purity. In crystallization, a diluent is added to the homogeneous solution that reduces the solubility of the product to the point that it “falls out” of solution and crystallizes. It is similar to precipitation but results in the formation of crystals rather than unordered aggregates. Crystallization can be used on a laboratory scale for determining protein structure, on on the industrial scale for antibody and therapeutic protein productions. Batch crystallizers are often used in industry because of their simplicity and inexpensiveness compared to continuous crystallization (Harrison et al., 2003).

==Membrane Separation==
Membrane separation takes advantage of the selective permeability of membranes; they allow certain particles to pass through and selectively stop other, generally unwanted, particles. The component that passes through is called the permeate and the component stream that is rejected is called the retentate or concentrate. The applicability of membranes comes from the fact that their selectivity is determined by their pore size, which can be controlled during the creation of the membranes. Additionally, Membrane processes do not require heat meaning they generally require less energy than conventional separations technology such as distillation and crystallization. Membrane separations processes are generally classified as microfiltration, ultrafiltration, or nanofiltration depending on the size of the particles to be filtered out.

[[File:membrane techs.png|frame|Figure. Cutoffs for different membrane categories]]

===Membrane Selection, Construction, and Flow Geometries===
Membrane permeability and selectivity are the two most important factors to consider when selecting a membrane. For gas separations, the permeation of the gas is usually facilitated by the gas dissolving in the membrane on one side and then evaporating on the permeate side. Therefore permeability depend largely on the solubility of components in the membrane.

The two most commonly utilized membrane configurations are hollow fiber and spiral wound. Hollow fiber is generally the most commonly utilized module for gas separations. These are formed by gluing the two ends of the hollow fiber to a resin forming a closure. The fibers are housed in a shell much like a heat exchanger. The feed flows past thousands of tubes with the permeate flowing into the hollow tubes and out the closure. The retentate then flows out of the shell not having gotten through the membrane.

[[File:hollow membrane.jpg|frame|Figure. Hollow fiber membrane module]]

Spiral wound membranes are created by sealing two membrane sheets back to back on three edges to form a sort of pocket. This fourth open edge is then attached to a porous tube which allows permeate to go through it. Several membrane pockets are attached to a single tube and wrapped around in a spiral.

[[File:spiral membrane.jpg|frame|Figure. Spiral wound membrane module]]

Flow geometry is usually either dead-end geometry or cross flow geometry. In dead end, the fluid flow is normal to the membrane surface while cross flow is parallel to the membrane surface. Dead end geometry is usually used with hollow fiber membranes while cross flow is used with spiral wound membranes. Each geometry has advantages and disadvantages. Dead end geometry is generally cheaper to set up and therefore has lower initial capital costs. However, it is very vulnerable to membrane fouling, which reduces the effectiveness of the membrane. This is usually the geometry set up for small scale lab experiments. The tangential flow devices are more cost and labor-intensive, but they are less susceptible to fouling due to the sweeping effects and high shear rates of the passing flow. Most commercial industrial membrane separations are done using spiral wound cross flow membrane modules.

===Applications===
====Food Industry====
Due to the fact that MD can be conducted at relatively low feed temperatures, it was successfully tested in many areas where high temperature applications lead to degradation of the process fluids especially in food processing. It was demonstrated that MD can be used for the concentration of milk, for the recovery of volatile aroma compounds from black currant juice, and for the concentration of many other types of juices including orange juice, mandarin juice, apple juice, sugarcane juice, etc.
====Reverse Osmosis====
Reverse osmosis is the most widely used membrane separation process. In this process, fresh water passes through the membrane while dissolved salts and other solids are rejected and stay in the concentrate. In this process, feed water is pressurized in order to overcome the osmotic potential difference between the salty retentate and the fresh water desired. These processes are generally run using spiral wound membrane cylinders using a cross flow setup.

===Membrane Model===
The two most important components when considering different membranes are the permeability, which will determine flux through the membrane, and selectivity, which will determine what passes through the membrane and how much. The flux through a membrane is defined as:
<math> M_i = \frac{P_i}{δ}(p_{i,f} - p_{i,p})</math>
Where Mi is the molar flux of component i, Pi is the permeability of the membrane for component i, δ is the membrane thickness, and pi,f and pi,p are the partial pressures of component i on the feed side and permeate side respectively.
The average flux across a long cylindrical membrane such as the spiral wound module is given by:
<math> \int_0^Lm \frac{M_i,dx}{L_m}</math>
Where Lm is the length of the cylinder and x is length in meters

Membrane selectivity of the ideal separation factor is given as the ratio of the permeability of one substance over another as shown:

<math> S_(i,j) = P_i/P_j </math>

where Sij is the selectivity of the membrane for component i over j.

==Cyclones==
Centrifugal Separators, more commonly know as Cyclones, are one of the most widely used gas-solid separators. They are typically used for particles that are 5μm or larger, but if agglomeration occurs they can sometimes separate particles as small as particles as small as .5μm. These cyclones can be designed for high efficiency separations or for high throughput (Towler, 2012).

[[File:cyclone.png|thumb|center|Reverse-flow Cyclone (Towler, 2008)|500x300px]]

Cyclones are popular in industry because of their low capital costs, low maintenance costs, low pressure drop, temperature and pressure limitations are only caused the materials, they work as dry equipment, and they don't take up much space (EPA).

They are not without their disadvantages. They have low collection efficiencies for small particles, they are unable to handle sticky materials and high efficiency units often have large pressure drops (EPA).

===Design===
Design Procedure
There are multiple design methods for cyclones. The deign method discussed below is the Stairmand method. This involved creating two designs one of which was for high throughput and the other which was designed for high efficiency. Using these base designs future cyclones can simply be estimated using scaling factors (Towler, 2012).

1. Select whether efficiency or performance is more important for your design

2. Determine the particle size distribution in your stream

3. Estimate the number of cyclones needed

3. Calculate the cyclone diameter

4. Calculate the scale up factor

5. Calculate the cyclone performance and overall efficiency

6. Calculate the cyclone pressure drop

7. Cost the system

====Particle Size====
The diameter of particles separated by cyclones are governed by this design equation. This equation is dependent upon the standards of the cyclone and how much you are trying to deviate from its standard operating conditions. Looking into the equation it shows that a larger internal cyclone diameter, higher flowrate, higher difference in density, and lower viscosity result in smaller particles being seperated. The whole term after d1 is known as a scale up factor. When designing a cyclone you would determine the flow rate, viscosity, particle size, and density difference would be known prior to using this design equation, so this would be used to optimize the diameter of the cyclone you were designing (Towler, 2012).

[[File:Towler 10.8.PNG|center|Reverse-flow Cyclone (Towler, 2012)|500x300px]]

d1=mean diameter of particle separated at the standard conditions, at the chosen
separating efficiency

d2= mean diameter of the particle separated in the proposed design, at the same
separating efficiency

Dc1 =diameter of the standard cyclone

Dc2 =diameter of proposed cyclone

Q1 =standard flow rate

Q2 =proposed flow rate

Δρ1 =solid-fluid density difference in standard conditions

Δρ2 =density difference, proposed design

μ 1=test fluid viscosity

μ 2=viscosity, proposed fluid.

[[File:Towler10.44.PNG||thumb|center|(a) Standard Cyclone Dimensions (b) High Gas Rate Cyclone (Towler, 2012)|500x300px]]

====Efficiency====
Efficiency is the percentage of total solids that are removed from the stream The graphs below were experimentally for the two standard designs Stairmand Method. This graph has the efficiency vs the particle size (Towler, 2012).

[[File:Towler10.45.PNG||thumb|center|Performance curves, standard conditions. (a) High-efficiency cyclone performance curves, standard conditions. (b) High gas-rate cyclone (Towler, 2008)|600x400px]]

If you have to scale the efficiency it can be estimated by the point where a horizontal line from the efficiency of the original particle size intersects the new particle size. A demonstration is on the graph below.

[[File:Towler 10.46.PNG||thumb|center|Scaled Performance Curve(Towler, 2008)|500x300px]]

====Pressure Drop====
One of the most important design factors in a cyclone is its pressure drop. The pressure drop occurs because entry and exit losses, kinetic energy losses, and friction losses that occur inside the cyclone. The pressure drop can be calculated using the following equation (Towler, 2012).

[[File:Towler 10.9.PNG|center|Reverse-flow Cyclone (Towler, 2008)|500x300px]]

The variables are defined as follows:
DP = cyclone pressure drop, millibars;
rf = gas density, kg/m3
;
u1 = inlet duct velocity

u2 = exit duct velocity

rt = radius of circle to which the center line of the inlet is tangential

re = radius of exit pipe

φ =factor from figure below

ψ = fc(As/A1)

fc = friction factor

As = surface area of cyclone exposed to the spinning fluid

A1 = area of inlet duct

[[File:Towler_10.47.PNG|thumb|center|Cyclone Pressure Drop Factor (Towler, 2008)|600x400px]]

====Cost Estimation====
The following values are represented in 2002 dollars. Cost is mostly a function of the amount of gas that need to be passed through the cyclone and the units are in standard meters cubed per second (sm3/sec).

The costs can be broken down into three categories:

Capital Costs: $4600-$7600 sm3/sec

Operations and Maintenance: $1500-$18,000 sm3/sec

Annualized Cost: $2800-$29000 sm3/sec

This all comes to $.47-$440 per metric ton of pollutant removed. The high variation in these costs have to do with smaller units being much less efficient per unit of pollutant removed. Furthermore the cost is also effected by the amount of particulates that are in the gas to begin with. For example a gas with 10 g/sm3 and air with 100sm g/sm3 would have similar pumping costs, but very different cost per unit of solids removed (EPA).

==Other Separation Processes==
===Extraction===
Liquid-liquid extraction is a process for components with overlapping boiling points and azeotropes. The process requires a solvent such that some of the components of the mixture are soluble, and then the components will be separated based on this solubility in the liquid. This process can operate at moderate temperatures and pressures, so is not very energy intensive. However, a distillation column is required to extract the solvent for recycle. More recently, supercritical fluids have replaced liquid solvents in some processes for L/L extraction, due to the solute’s ability to more rapidly diffuse through them. The issue with these fluids, however, is that they must be operated at extremely high pressures and temperatures, increasing both capital and operating expenses of the process (Peters & Timmerhaus, 2003).

===Crystallization===
This process recovers solutes that have been dissolved in solution. The resulting product is in the solid phase. Depending on the material properties of the solute and solvent, the solute is recovered by precipitation after cooling, removal of solvent, or adding precipitating agents. Crystallizers are designed based on phase equilibria, solubilities, rates and amounts of nuclei generated, and rates of crystal growth. Every crystallization process is a unique system, so plant evaluation is usually required before complete implementation. Crystallization can be performed in both batch and continuous processes, and design features can control crystal size to an extent (Peters & Timmerhaus, 2003).

===Membrane Separation===
This separations process uses selectively permeable membranes to separate components in a mixture. Typically, one of the components will freely pass through the barrier while the other components will not. The stream that passes through the membrane is the permeate and the stream that does not pass is the retentate. The driving force behind this separation is a pressure gradient. Membrane separation is beneficial because it can separate mixtures at the molecular and small particle level. Furthermore, there is no phase change required so the energy input is low. Limitations of this process include achieving high product purity, incompatibility with certain stream components, low operating temperature, and low flow rates. Although membrane separation is generally not scaled up, examples of scaled-up membrane separation include seawater desalination and hydrogen recovery (Peters & Timmerhaus, 2003).

===Adsorption===
Adsorption involves an adsorbent and adsorbate. The adsorbent is typically a solid, and will typically separate the adsorbate from the stream. This process usually includes a desorption step that regenerates the adsorbent for further use. Raising the temperature or increasing the concentration of the adsorbate can reverse the adsorption process. Although the recycle of the adsorbent is a very economic design feature, the downside of this step is that it results in a cyclic process, which introduces complexity to the overall process. Industrial applications of this process are for bulk separations and gas purification. The adsorption/desorption process in these situations involves a large amount of heat transfer, which design engineers must take into account when sizing and selecting equipment material (Peters & Timmerhaus, 2003).

===External Field/Gradient Separation===
These separations use external force fields or temperature gradients to separate responsive molecules or ions. The use of these processes is fairly limited to a few specialized industrial applications (Peters & Timmerhaus, 2003).

===Settling and Sedimentation===
In settling processes, solid particles or liquid drops are separated from a stream by gravity. The stream can be in either the liquid or gas phase. For vapor-liquid mixtures, flash drums are generally used to separate the mixture. The velocity of the vapor must be less than the settling velocity of the liquid drops for this separation to occur. For liquid-liquid separation, the horizontal velocity of the fluid must be low enough to allow the low-density droplets to rise to the interface and the high-density droplets to move away from the interface and coalesce. In sedimentation, the result of the process is a more concentrated slurry. Typically a flocculating agent is used to aid in the settling process. One way to perform this separation is to use a cone-shaped tank with a slowly revolving rake that scrapes and moves the thickened slurry to the center of the cone for removal (Peters & Timmerhaus, 2003).

[[File:Example.jpg]]

====Clarifiers====
[[File:Circular_Clarifier.png|300px|thumb|bottom|Figure 9: Circular clarifier with some components labelled.]] [[File:Rectangular_Clarifier.png|300px|thumb|bottom|Figure 10: Rectangular clarifier with some components labelled.]]

Clarifiers are one of the methods used for the continuous removal of particulate solids from liquids through sedimentation by gravity. Applications include process water pretreatment, waste water treatment, and drinking water purification. Historically, clarifiers were originally developed to limit nutrient input into surface water due to fear of eutrophication. Today, they have a number of uses, particularly in wastewater treatment processes, metal removal, disinfection, and membrane pretreatment. The process helps removed dissolved solids, silt, and undesirable metals from the water, making it more suitable for downstream processes as well as human consumption (Wilson, 2005).

Clarifiers are typically used in conjunction with coagulation or flocculation agents, which promote dissolved particles to join into clumps and settle out of solution (Towler and Sinnot, 2012). Clarifiers typically consist of a large circular tank with a rotating rake at the base which scrapes settled solids towards the center. In the case of a rectangular clarifier, they are scraped to one side. Diagrams of both are represented in figures 9 and 10, respectively (NMED Surface Water Quality Bureau, 2015). Separated solids are allowed to settle to the bottom of the tank as a sludge, whereupon they are collected by the rake and disposed of properly. In the case of floating contaminants, it is possible for the clarifier to include a skimmer as well.

Clarifier efficiency varies with certain factors, including the settling characteristics of solids removed and the surface overflow rate of the tank. Clarifier efficiency can be found using the following relation:

<math>\begin{align}
E_{TSS} &= E_{TSSmax}\left ( 1 - e^\frac{\lambda}{SOR} \right )
\end{align}</math>

where <math>E_{TSS}</math> is the efficiency of total suspended solids (TSS) removal, <math>E_{TSSmax}</math> is the maximum possible efficiency, <math>\lambda \left [\frac{m}{d} \right ]</math> is the settling constant, and <math>SOR \left [\frac{m^3}{m^2 d} \right ]</math> is the surface overflow rate. The effect of flocculation chemicals on TSS can be seen in figure 11. However, it should be noted that chemical addition will increase sludge quantity and may have an adverse effect on plant aesthetics, which increases maintenance costs (Wilson, 2005).

[[File:Chem_Addition.png|200px|thumb|bottom|Figure 11: The effect of flocculating agents on total suspended solids removal in clarifiers.]]

=====Lamella Clarifiers=====

Lamella clarifiers use inclined plates in order to maximize the settling area for solids. Solids continue to settle into a hopper at the bottom of the tank while clarified water exits up through the inclined plates. This allows for the design of a smaller tank, which leads to large savings in capital costs. A lamella clarifier is pictured in figure 12.

[[File:Lamella_Clarifier.png|300px|thumb|bottom|Figure 12: A lamella clarifier with components labeled.]]

Typically, inclined plates are installed at an angle of 45 to 60 degrees and spaced 40 to 120 mm apart, which increases effective settling surface area by a factor of 6 to 12 compared to traditional clarifiers. For effective use, it is recommended that the Reynolds number be below 2000, Froude number higher than 10-5,and detention time be longer than 3 to 5 minutes. For this implementation, the equations are as follows:

<math>\begin{align}
N_{Re} &= \frac{VR}{\nu} \\
N_{Fr} &= \frac{V^2}{Rg}
\end{align}</math>

where <math>R</math> refers to the hydraulic radius, which is the cross-sectional area of the lamella, <math>V</math> is the liquid velocity, <math>\nu</math> is the kinematic viscosity, and <math>g</math> is the gravitational constant (Wilson, 2005).

=====Advantages=====

Clarifiers offer a proven, relatively inexpensive solution for solids removal. The chemical coagulants used are cheap and provide a low operating cost as well as simple maintenance. Construction is typically simple, leading to low capital costs and equipment that is easy to accommodate and maintain. Their design is also flexible, with various options such as skimmers and scrapers offering increased removal efficiency (Wilson, 2005). Operation of clarifier tanks also has lower energy requirements than membrane filtration for solids removal, given that most of the separation is aided by gravity. Water exiting clarifier units has a silt density index (SDI) averaging 4.0, which is low enough for further membrane treatment such as reverse osmosis (Prihasto, 2009).

=====Disadvantages=====

Clarifiers necessitate low turbulence to prevent resuspension of solids. This essentially requires a low entrance velocity, which can limit the production rate of certain processes or call for more clarifier units, which would drive up costs. Furthermore, clarifiers require frequent cleaning before sludge becomes too difficult to remove and reduces effectiveness. In the case of lamella clarifiers, sludge buildup on the inclined plates results in uneven flow distribution which could harm efficiency (US EPA, 2003). For this reason, maintenance requirements for lamella clarifiers are higher, but they can be reduced through the implementation of removable plates (Wilson, 2005). Clarifiers also only remove solids, so pH will not be affected, leading to the need for further pH adjustment (NMED Surface Water Quality Bureau, 2015).

=====Clarifier Design Calculations and Typical Design Values=====

======Detention Time======

Detention time (DT) is the time is takes for a unit of water to travel from the inlet of the clarifier unit to the outlet. During typical operations, the design value for this is 2 to 3 hours.

<math>\begin{align}
DT &= \frac{Tank\ Volume}{Influent\ Rate}
\end{align}</math>

======Surface Overflow Rate======

Surface overflow rate (SOR) measures the flow into the clarifier per square foot of surface area. Typical design values are 400 to 800 gal/day/sq. ft.

<math>\begin{align}
SOR &= \frac{Volumetric\ Flow\ Rate}{Surface\ Area}
\end{align}</math>

======Weir Overflow Rate======

Weir overflow rate (WOR) describes the flow in gallons per day per linear foot of weir. Typical values are 10,000 gal/day/ft.

<math>\begin{align}
WOR &= \frac{Volumetric\ Flow\ Rate}{Weir\ Length}
\end{align}</math>

======Solids Loading Rate======

Solids loading rate (SLR) describes the mass of solids in the clarifier influent per square foot of surface area. This value should not exceed 30 lbs/day/sq. ft.

<math>\begin{align}
SLR &= \frac{Solids\ Mass\ Flow\ Rate}{Surface\ Area}
\end{align}</math>

===Flotation===
Flotation is a process designed for specific solid-solid mixtures. It works by generating gas bubbles in a liquid that attach to selected solid particle. Afterwards, the particles rise to the liquid surface where they are removed by an overflow weir or mechanical scraper. The separation depends on the surface properties of the particles and its preference to attach to the gas bubbles. To meet the necessary requirements of the flotation process, a number of additives can be used to control things like the pH of the liquid-solid mixture, the activity of the solid surface, and the froth that can assist in separation. The bubbles can be produced by gaseous dispersion, dissolution, or electrolysis of the liquid (Peters & Timmerhaus, 2003).

===Centrifugation===
This process is similar to external field separation in that an external force field is applied to separate a mixture. When gravity separation is too slow due to particle densities, particle size, settling velocity, or the formation of an emulsion, centrifugation is commonly used. Centrifugal force increases the total force acting on the particle and results in faster separation times. This process is generally used to separate solids from liquids, however it can also be used to separate two liquids with very different densities (Peters & Timmerhaus, 2003).

===Drying===
Drying is performed to remove liquid from a liquid-solid mixture and produce a dry solid. Water is most often the liquid removed, but organic liquids are removed from solids on occasion as well. The heat required to vaporize the liquid is usually obtained by a series of gas-solid contacting devices. Feed condition and temperature sensitivity of the solid dictate the type of contacting device that is used. There are two groups of dryers that differ by the dependence of either mechanical means or fluid motion for gas solid contact. Another feature of dryers is to use either direct (hot gas) or indirect (conductive surface) heating (Peters & Timmerhaus, 2003).

===Evaporation===
Evaporators separate solvents from a solution by evaporation. The difference between evaporation and distillation is that evaporation requires the solute be nonvolatile. Because of this, a high separation can be achieved with one stage. Evaporators are essentially reboilers, so evaporation is a very energy-intensive process with a high thermal economy (Peters & Timmerhaus, 2003).

===Filtration===
Filtration is a process that separates a mixture of solid in a liquid or gas by passing the mixture through a porous medium in which the particles do not pass. Filtration is done by either cake filtration (particles found on the surface of the filter) or depth filtration (particles found within the filter). Cake filtration is generally performed with a cloth as the filtration medium (Peters & Timmerhaus, 2003).

==Conclusion==
Separation is a key part of most chemical processes, and there is a great variety of techniques to perform separation of compounds based on size, volatility, charge, and many other features. A common technique with which the process engineer should be familiar is distillation, but he or she should also be aware of the other available options. Some techniques may be less expensive, less energy-intensive, or more effective than distillation, depending on the specific separation problem. Therefore, the separation strategy should be carefully considered.

==References==
Air Control Technology Fact Sheet. Enviornmental Technology Fact Sheet. http://www3.epa.gov/ttncatc1/dir1/fcyclon.pdf

Belter PA, Cussler EL, Hu WS. Bioseparations: Downstream Processing for BIotechnology. New York: John Wiley; 1998.

Biegler LT, Grossmann IE, Westerberg AW. Systematic Methods of Chemical Process Design. Upper Saddle River: Prentice Hall; 1997.

Danckwerts P (1965) The Absorption of Gases in Liquids. Pure and Applied Chemistry UK 10:625-642.

Development Document for the Final Effluent Limitations Guidelines and Standards for the Metal Products and Machinery Point Source Category (Report). US Environmental Protection Agency. 2003.

Erwin, D. Industrial Chemical Process Design. New York: McGraw Hill, Professional Engineering; 2002.

Hage, David S. 1999Affinity Chromatography: A Review of Clinical Applications. Clinical Chemistry 45(5): 593–615.

Harrison RG, Todd P, Rudge SR, Petrides, DP. Bioseparations Science and Engineering. New York: Oxford University Press; 2003.

Hage, David S. 1999Affinity Chromatography: A Review of Clinical Applications. Clinical Chemistry 45(5): 593–615.

HisPur Cobalt Resin - Thermo Fisher Scientific N.d. https://www.thermofisher.com/order/catalog/product/89964, accessed February 21, 2016.

Lamella Plate Clarifier. Hydro International Web site. Available at: http://www.hydro-int.com/uk/products/lamella-plate-clarifier?s=0&r=uk. Accessed February 2, 2016.

Lean Oil Absorption. PetroGas Systems Web site. Available at: http://petrogassystems.com/technology/natural-gas-processing-and-dew-point-control/lean-oil-absorption. Accessed February 19, 2014.

Lichty, Jordan J., Joshua L. Malecki, Heather D. Agnew, Daniel J. Michelson-Horowitz, and Song Tan 2005Comparison of Affinity Tags for Protein Purification. Protein Expression and Purification 41(1): 98–105.

Merichem Gas Technologies. ®LO-CAT PROCESS available at http://www.merichem.com/images/casestudies/Desulfurization.pdf Accessed 6 Feb. 2015.

Miller L.N. & Zawacki T.S. , US 4080424, "Process for acid gas removal from gaseous mixtures", issued 21 Mar 1978, assigned to Institute of Gas Technology

NMED Surface Water Quality Bureau, New Mexico Water Systems Operator Certification Study Manual, New Mexico Environment Department, 2015.

Peters MS, Timmerhaus KD. Plant Design and Economics for Chemical Engineers. 5th ed. New York: McGraw Hill; 2003.

Prihasto, N; Lui, Q; Kim, S. Pre-treatment strategies for seawater desalination by reverse osmosis system. 2009; 249(1): 308-316. doi:10.1016/j.desal.2008.09.010

Schmidt Eberhard (2012) Waste Gases, Separation and Purification. Ullman’s Encyclopedia of Industrial Chemistry Germany 2:174-181.

Seider, W.D., Seader, J.D., and Lewin, D.R. (2004). ''Process Design Principles: Synthesis, Analysis, and Evaluation.'' New York: Wiley.

Stripping Column. Alfa Laval Web site. Available at: http://www.alfalaval.com/solution-finder/products/soft-column/Documents/Stripping%20Column.pdf. Accessed February 19, 2014.

Towler G, Sinnott R. Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013.

Turton, R.T., Bailie, R.C., Whiting, W.B., and Shaewitz, J.A. (2003). ''Analysis, Synthesis, and Design of Chemical Processes'' Upper Saddle River: Prentice-Hall.

Wankat, P.C. (2012). ''Separation Process Engineering.'' Upper Saddle River: Prentice-Hall.

Wilson, T.E., Clarifier Design, 2nd Ed., McGraw-Hill: New York, 2005.

Separation processes

2016-02-22T01:33:39Z

Bts511: /* Cost Estimation */

Authors: Nick Pinkerton, [2014] Karen Schmidt, [2014] James Xamplas, [2014] Emm Fulk, [2015] and Erik Zuehlke, [2015] John Dombrowski [2016] , Brett Sleyster [2016] , Robert Cignoni [2016] , and Osman Jamil [2016] 

Stewards: David Chen, Jian Gong, and Fengqi You

Date Presented: February 9, 2014 /Date Revised: February 1, 2014

 

==Introduction==
Essentially all chemical processes require the presence of a separation stage. Most chemical plants comprise of a reactor surrounded by many separators. Separators have a countless number of jobs inside of a chemical plant. A separator can process raw materials prior to the reaction, remove incondensable gases, remove undesired side products, purify a product stream, recycle materials back into the process, and many other jobs that are essential to the process.

Chemical engineers must understand the science of separation and the variety of ways that separation can take place. There are many ways to perform a separation some of these including: distillation, absorption, stripping, and extraction. The science of separation revolves around the presence of two phases that are in contact and equilibrium (Wankat, 2012).

[[File:Sepmeth.JPG|frame|Figure 1. Separation methods by property]]

==Theory==
===Vapor-Liquid Equilibrium===
Separation processes are based on the theory of vapor-liquid equilibrium. This theory states that streams leaving a stage in a separation process are in equilibrium with one another. The idea of equilibrium revolves around the idea that when there is vapor and liquid in contact with one another they are in constantly vaporizing and condensing. Different components in the mixture will condense and vaporize at different rates. There are three types of equilibrium conditions that can be subdivided into thermal, mechanical and chemical potential categories. These separate equilibrium states are given as:

<math>T_{liquid} = T_{vapor}</math>

<math>p_{liquid} = p_{vapor}</math>

<math>chemical potential_{liquid} = chemical potential_{vapor}</math>

==Distillation==
===Flash Distillation===
Flash Distillation is one of the simpler separation processes to be employed in a chemical plant. The main premise of flash distillation is that a portion of a liquid feed stream vaporizes in a flash chamber or a vapor feed condenses. Vapor-liquid equilibrium will cause the vapor phase and the liquid phase to have different compositions. The more volatile component of the mixture will compose of a larger portion of the vapor. This simple separation is easy to manufacture but does not result in large degrees of separation.

Flash distillation requires a feed stream that is pressurized and heated and then passed through a valve into a flash drum. The large pressure drop across the valve will result in a partial vaporization of the fluid. Vapor will be removed overhead from the flash drum while the remaining liquid will collect at the bottom of the drum and be removed. Most flash drums will contain an entrainment eliminator which is a screen that prevents liquid from being carried into the vapor effluent. Figure 2 shows a simple overview of the flash distillation process. As shown, there is a heater that flows into a let-down valve where the two-phase flow begins. Variables y and x are the mole fractions of the more volatile component in the vapor and liquid effluents, respectively.

[[File:Flash.gif|center|frame|Figure 2. Flash Distillation Flow Diagram]]

===Column Distillation===
Distillation columns are the most widely used separation technique used in the chemical industry, accounting for approximately 90% of all separations (Wankat, 2012). Distillations in columns consist of multiple trays that each act at their own equilibrium conditions. Large columns are able to perform complete separations of binary mixtures as well as more complex multi-component mixtures.

[[File:column.jpg|250px|center|]]
===Stages===
Columns are separated into stages by the presence of trays. These trays allow for vapor-liquid contact and equilibrium to occur. Typically, the more stages in a column, the larger separation that can be achieved. There are many different types of trays that can be used in a column.
====Sieve Trays====
The simplest and least expensive tray type is the sieve tray which is a sheet of metal with holes punched into it to allow vapor flow. Sieve trays can have different hole patterns and sizes that will affect the tray efficiency and flow rates.

[[File:sieve.jpg|200px|center|]]

====Sieve Tray Design Procedure====

The design of these plates is done through a trial-and-error process. Most commercial process simulations (such as HYSYS) have default tray designs, and automatically specify dimensions. However, these dimensions selected or calculated by the simulations may not give the best performance for your system, so it is valuable to understand how to design the sieve trays and how specific parameters may affect performance. Hand calculations using the following methods can be used to guide the simulation programs to better design. This section will use sample data to work through an example of the process. The following is a general list of steps for designing a sieve plate:

=====1. Calculate the maximum and minimum vapor and liquid flow rates for the turndown ratio required.=====
This data can be collected from a McCabe-Thiele diagram and/or from process simulation data.

Data from McCabe Thiele diagram, for example: 
Number of stages = 10 
Slope of top operating line = 0.185 
Slope of bottom operating line = 1.43 
Top composition = 98.8 mol% acetone 
Bottom composition = 4 mol% acetone (a) 
Minimum reflux ratio = 0.31 

Collect data rom mass balances: 
MW of feed = (0.6 kmol a/kmol)*(58 kg/kmol a) + (0.4 kmol acetic acid/kmol)*(60 kg/kmol aa) = 59 kg/kmol 
Feed rate = 100 kmol/h 
Using material balances... D = 57.3 kmol/h ; B = 42.7 kmol/h 
Vapor rate, V = D(1+R) = 57.3 kmol/h (1+.31) = 75.1 kmol/h 
Below feed line, slope of BOL = 1.43 = L'm/V'm (liquid rate/vapor rate)

=====2. Collect or estimate the system physical properties.=====
Here it is important to know information about both the top and bottom of the column. Useful information includes temperature, pressure, column pressure drop (a common assumption is 100 mmH2O per plate), densities, molecular weights, surface tensions, and number of stages (which can be estimated from the McCabe-Thiele diagram).

=====3. Select a Trial Plate Spacing=====
The plate spacing will depend on the column diameter and operating conditions. Plate spacings from 0.15 m to 1.0 m are typically used. The smaller the diameter, the smaller the spacing. Small columns will use close spacing. Columns with diameters above 1.0 m, plate spacings of 0.3 m to 0.6 m are normally used. A good initial estimate is 0.5 m.
[[File:trayspacing.jpg|400px|center|]]

=====4. Estimate the column diameter, based on flooding considerations.=====
Vapor and liquid flow rates will vary along the column, so plate design needs to be considered both above and below the feed. Using plate spacing and FLV (which is the square root of the ratio of the liquid to vapor flow rates), you can obtain the value of K from the plot.

[[File:floodingplot.jpg|400px|center|frame|Figure. Plate Spacing]]
There is a range of vapor and liquid flow rates in which the column needs to be operated. Too low or too high of rates can result in various inefficiencies in the column operation, as shown in the figure below. For example, if the vapor rate is too high, flooding will occur. However, it is not safe to operate on the flooding line. Instead, columns are typically designed for 80% of flooding at the maximum flow rate.
[[File:vap_rate_vs_liq_rate.jpg|400px|center|frame|Figure. Tray behavior]]

=====5. Decide the liquid flow arrangement.=====
Common flow arrangements are single pass (cross flow), double pass, and reverse flow. Using conditions at the bottom of the column, calculate the max volumetric flow rate. Use this flow rate and the column diameter to determine the preferred flow arrangement from the chart below.
[[File:Liquidflow.jpg|200px|center|]]

=====6. Make a trial plate layout: downcomer area, active area, hole area, hole size, weir height.=====
Standard sizes for trays -- and good assumptions for the first iteration -- are: weir height, hw = 50mm ; hole diameter, Dh = 5mm ; plate thickness, tpl = 5mm. From the graph below, the ratio of downcomer area (Ad) to column cross-sectional area (Ac) can be determined from the ratio of weir length (lw) to column diameter (Dc) and vice versa.
[[File:platelayout.jpg|200px|center|frame|Figure. Plate Dimensions]]

=====7. Check the weeping rate=====
Compare the actual vapor velocity to the minimum vapor velocity -- if velocity is too low fluid will "weep" through the tray holes. If the weeping rate is unsatisfactory, return to step 6 and choose different values for the plate layout dimensions. From the chart in step 4, it can be seen that there is a minimum vapor flow rate below which the liquid "weeps" from the tray above.

For the remaining steps in this design process, it is recommended to check your assumptions after each step and revise them as necessary in order to maintain operation in the "sweet spot" of the vapor rate vs. liquid rate plot. Additional iterations may be required as you move through the procedure.

Calculate the maximum liquid flow rate. Calculate the minimum liquid flow rate at 70% turndown (recommended). Calculate the height over the weir as
<math>h_o=750[\frac{L_w}{p_Ll_w}]^\frac{2}{3}</math>

=====8. Check the plate pressure drop=====
<dfn>If the pressure drop calculated here is too high, return to step 6.</dfn>

Proceed to step 9 if the pressure drop assumption is valid.
=====9. Check the downcomer backup. =====
<dfn>If the downcomer backup is too high, return to step 6 or 3.</dfn>

The plate spacing affects the amount of fluid in the downcomer. Calculate the level in the downcomer and the residence time of the fluid to see if the values are valid. Note that residence times greater than 3 seconds are acceptable.

Proceed to step 10 if residence time is acceptable.
=====10. Decide plate layout details.=====
Determine calming zones, the unperforated areas at the inlet and outlet sides of the plate. The width of each zone is usually made the same. Recommended values are: below 1.5 m diameter, 75 mm; above, 100 mm. The unperforated area can be calculated from plate geometry. Also check the hole pitch, or the distance between hole centers. It should not be less than 2.0 hole diameters. A normal range is between 2.5 and 4.0 hole diameters. The shape must also be specified. Square and equilateral triangle holes are used.

=====11. Recalculate the percentage flooding based on the chosen column diameter.=====
An assumption of 80% flooding was chosen so that operation would occur in the "sweet spot." This assumption must be checked by calculating the flooding percentage for a given column diameter.
uv = (max volumetric flow rate)/(net area)

<math>%flooding = \frac{u_v}{u_f}</math>

If the hole pitch is unsatisfactory, return to step 6.
=====12. Check entrainment=====
''If too high, return to step 4'' Use the graph below to determine entrainment from FLV.
[[File:entrainment.jpg|400px|center|]]

The value for fractional entrainment can be used to re-estimate the column efficiency, and reevaluate the number of trays needed. Can return to step 1 for more accurate estimates.

=====13. Optimize design.=====
After returning to step 1 to reevaluate the number of trays, it is valuable to repeat steps 2 through 12 to find the smallest diameter and plate spacing acceptable at the lowest cost.

=====14. Finalize the design.=====
Optional: draw up the plate specification and sketch the layout of the plate.

====Bubble-Cap Trays====
Bubble-cap trays consist of a weir around each hole in the tray which is covered with a cap that has holes or slots to allow vapor passage. Entrainment is about three times larger than a sieve tray. Bubble-cap trays require larger tray spacing than sieve tray design. Bubble-cap trays have been known to have problems with coking, polymer formation, or high fouling mixtures. Recently, very few new bubble-cap columns are being built due to the expense and marginal benefits. However, engineers will likely encounter bubble-cap columns still currently in operation.

====Flow Patterns====
Cross flow columns are the most common pattern for distillation columns. For liquid flows between 50 and 500 Gal/min, a cross flow column is appropriate. When liquid flow is increased above 500 Gal/min, an engineer should consider designing a double pass or multi-pass column. This will reduce the liquid gradient on the tray and reduce the downcomer loading (Wankat, 2012).

===Column Sizing===
Column height will be dependent on the amount of trays required and the spacing between the trays. Normally, tray spacing of 0.15 m to 1 m is used. For columns, above 1 meter in diameter, 0.5 m can be used as an initial estimate.

Column diameter is influenced by the vapor flow rate in the column. The trays can not have excess liquid entrainment or high pressure drops; therefore, vapor velocity in the column must be maintained at a reasonable level.

An equation based on the Souders and Brown equation can be used as an estimate for the max allowable superficial vapor velocity,

<math>\hat u_v = (-0.171l_t^2 + 0.27l_t - 0.047){\frac{\rho_L - \rho_v}{\rho_v}}^{1/2}</math>

where <math>l_t</math> is the plate spacing in meters, <math>\rho_L</math> is the density of the liquid stream, and <math>\rho_V</math> is the density of the vapor stream.

Column diameter, <math>D_c</math>, can then be estimated using the relation,

<math>D_c = \sqrt{\frac{4\hat{V_w}}{\pi\rho_v\hat{u_v}}}</math>

where <math>\hat{V_w}</math> is the maximum vapor rate in kg/s (Towler et al., 2013).

===Distillation Applications===

Distillation is a process that can be implemented in various scales. There is both laboratory scaled distillation as well as very large industrial distillation. Other applications for distillation include food/alcohol processing and herb distillation for the perfume and medical industries. Typically laboratory scaled distillation occurs in batches whereas industrial distillation (e.g. fractional distillation of crude oil) occurs continuous with a constant distillate and bottom effluent streams.

Some applications of distillation are concerned the top stream only, some the bottom stream only and others both streams can be used for future products. In alcohol distillation for example, the water that is separated from the ethanol/water binary solution is discarded as waste water. In fractional distillation of crude oils, the heavy hydrocarbons at the bottom of the column are collected and sold along with the light hydrocarbons that appear in higher side draws (Wankat, 2012).

===Example Case: Ideal Distillation===

Assume an equimolar mixture flowing at 10 mol/s of 20 mol% n-pentane, 30 mol% n-hexane, and 50 mol% n-heptane. Separate the mixture into 3 products: 99% pure n-pentane, 99% pure n-hexane, 99% n-heptane. Assume the feed and products are all liquids at the bubble points. There are two process alternatives to consider in this example. The direct sequence removes the most volatile species, pentane, in the first column, and then separates hexane and heptane in the second column. The indirect sequence separates the heaviest product, heptane, and then separates pentane from hexane in the second column. This example will consider the direct sequence. Next, we must decide if these species exhibit fairly ideal behavior during distillation. Since the n-alkanes have very similar properties, it is safe to assume they will display close to ideal behavior. The next step is to look up the boiling points of the 3 species. In this case, the normal boiling points of pentane, hexane, and heptane are 309 K, 342 K, and 372 K, respectively. Also, it is a good idea to look up relative volatilites, to further verify near-ideality of the mixture, but also to obtain the information necessary for the Underwood method, which we will employ to obtain a solution. The next step is to write out material balances based on molar flows and the design specifications. They go as follows:

<math>\mu_I(nC5) + \mu_{II}(nC5) = 2 mol/s</math>

<math>\mu_I(nC6) + \mu_{II}(nC6) + \mu_{III}(nC6) = 3 mol/s</math>

<math>\mu_{II}(nC7) + \mu_{III}(nC7) = 5 mol/s</math>

<math>\mu_I(nC5) = 99\mu_I(nC6)</math>

<math>\mu_{II}(nC5) = (5/990)\mu_{II}(nC6)</math>

<math>\mu_{II}(nC7) = (5/990)\mu_{II}(nC6)</math>

<math>\mu_{III}(nC7) = 99\mu_{III}(nC7)</math>

where <math>\mu</math> represents the molar flow, and the subscript represents the product stream.

Solving this system of equations:

<math>\mu_I(nC5) = 1.985\ mol/s</math>

<math>\mu_{II}(nC5) = 0.015\ mol/s</math>

<math>\mu_I(nC6) = 0.020\ mol/s</math>

<math>\mu_{II}(nC6) = 2.930\ mol/s</math>

<math>\mu_{III}(nC6) = 0.050\ mol/s</math>

<math>\mu_{II}(nC7) = 0.015\ mol/s</math>

<math>\mu_{III}(nC7) = 4.985\ mol/s</math>

At this point we have enough information to use Underwood's method to estimate the minimum vapor flows in the column. The following three equations are used in Underwood's method:

<math>\sum_i \frac{\alpha_{ik}}{\alpha_{ik}-\phi}f_i = (1-q)F</math>

<math>(R_{min}+1)D = \sum_i \frac{\alpha_{ik}}{\alpha_{ik}-\phi}d_i = V_{min}</math>

<math>\bar R_{min}B = -\sum_i \frac{\alpha_{ik}}{\alpha_{ik}-\phi}b_i = \bar V_{min}</math>

where <math>\alpha_{ik}</math> is the relative volatility of species <math>i</math> to species <math>k</math>, <math>f_i</math> the molar flow of species <math>i</math> in the feed, <math>q</math> the fraction of the feed that joins the liquid stream at the feed tray, <math>F</math> the total molar flow of the feed, <math>D</math> the molar flow of the distillate, <math>R_{min}</math> the minimum reflux ratio <math>(=L_{min}/D)</math>, <math>d_i</math> the molar flow of species <math>i</math> in the distillate, <math>V_{min}</math> the minimum vapor flow possible in the top section of the column to accomplish the desired separation, <math>\bar R_{min}</math> the minimum reboil ratio <math>(=\bar V_{min}/B)</math>, <math>b_i</math> the molar flow of species <math>i</math> in the bottoms product, and <math>\bar V_{min}</math> the minimum vapor flow in the bottom section of the column. The final variable, <math>\phi</math>, will be solved for using the first Underwood equation, and it's value will be decided based on the relative volatilities of the key components in the column.

So, after solving the first Underwood equation, we get two values for <math>\phi</math>, 3.806 and 1.462. Because 3.806 is between the relative volatilities of the key components, we will substitute that value for <math>\phi</math> into the second Underwood equation. Doing so for both columns gives <math>V_{min} = 6.4\ mol/s</math> for the first column and <math>V_{min} = 8.9\ mol/s</math> for the second column, for a total minimum vapor flow of 15.3 mol/s. The process would then be repeated for the indirect sequence, and the decision for which process to use would be justified by the process with the overall minimum vapor flow (Biegler et al., 1997).

==Absorption==
===Description of Absorption===
Another separation process used in industry is absorption, which is used to remove a solute from a gas stream. It accomplishes this by contacting the gas mixture with a liquid solvent that readily absorbs the undesirable components from the gas stream, purifying the gas stream. This separation process is determined by the inputs of the liquid flow rate, temperature, and pressure.

The absorption factor, which can be determined mathematically, determines how readily a component will absorb in the liquid phase. The absorption factor of component i is

<math>A_i=L/K_iV</math>

where <math>L</math> is the liquid flow rate entering the column, <math>V</math> is the vapor flow rate entering the column, and <math>K_i</math> is the vapor/liquid equilibrium ratio for component i (Peters & Timmerhaus, 2003). Higher absorption factors result in higher absorptivity into the liquid and a decrease in the number of trays required for separation, however a diminishing return occurs after the absorption factor is greater than 2.0. An absorption factor of 1.4 is most commonly used.

In general absorption can be seperated into two overarching categories, physical and chemical absorption. In physical absorption, the unwanted solute in the gas is absorbed into the liquid phase because solubility of the component is higher in the liquid phase than the gas phase. In chemical absorption the solute is removed from the gas via a reaction with the solvent, this reacted product is then transported into the liquid phase (Danckwerts 1965). There are two types of chemical absorption reversible and irreversible. Generally reversible chemical absorption is preferred as the solvent can be put through a stripper and regenerated so it can be recycled back to the absorption process (Wankat, 2012).

===Absorption Apparatus===

There are five major apparatus used for absorption in industrial application. These five pieces of equipment are spray absorbers (or towers), ejector (venturi) scrubbers, packed columns, trayed columns, and film absorbers (Schmidt, 2012).

==== Spray Tower vs Ejector Scrubber ====

In both '''spray tower''' and the '''ejector scrubber''' nozzles are employed to produce small solvent droplets. These small droplets increase the surface area of the liquid to gas contact allowing for the maximum amount of mass transfer to occur between the gas mixture and the liquid. The major difference between the two nozzle equipment designs is the configuration and type of nozzles. In the ejector scrubber shown in Figure 3 there is a single nozzle that is generally a higher pressure spray nozzle that produces finer solvent drops allowing for an even greater amount of mass transfer enabling better physical absorption (Schmidt, 2012).
[[File:Ejectorventuri.jpg|thumb|200px|center|Figure 3. Ejector Scrubber (US EPA, 2006)]]
'''Spray towers''' on the other hand generally have many nozzle at different heights where the liquid solvent will be sprayed out of to contact the gas running through the tower. This design is used in order to ensure the gas contacts the liquid as throughout the tower. These nozzles are lower pressure than a ejector scrubbers nozzle and thus physical mixing is worse in this configuration. Since physical mixing is generally worse in this configuration it is usually used in conjunction with a chemical absorption process. The other major difference between the ejector scrubber and the spray tower is that gas and liquid flow is cocurrent in the former while it is countercurrent in a spray tower. A spray tower absorber is shown below in Figure 4 (Schmidt, 2012).
[[File:SparyTowerAbsorber.jpg|thumb|200px|center|Figure 4. Spray Tower Absorber (US EPA, 2006)]]

==== Tower Type Absorption Apparatus ====
'''Packed column absorbers''' and '''tray column absorbers''' have very high efficiencies for the removal of an unwanted solute in the gas stream. The major disadvantage a trayed column has when compared to a packed column is the pressure drop. The pressure drop in a packed column is generally very low, whereas in between each tray of a trayed column pressure drop can be quite large. However the advantages inherent to trayed columns become clear when one needs the solvent to have a high concentration of the component to be removed from the gas stream. This is most important in the case where there is a very low concentration of the component in the gas stream and the specification states the solvent must contain a high concentration of that component. In this case the flow rate of the solvent may not be high enough for a packed column, however in a trayed column the solvent flow rate can be near zero for operation (Schmidt, 2012). Packed and trayed column internals are very similar to the setups found in the respective distillation columns.

For a '''trayed column''' the plate efficiency can be calculated using O'Connell's Correlation which invovles the Henry's Law constant, total system pressure, and solvent viscosity at the operating temperature (Towler & Sinnott, 2013).

<math>x=0.062*\frac{\rho_s*P}{\mu_s*H*M_s}</math>

where
<math>x</math> is the tray efficiency,
<math>\rho_s</math> is the density of the solvent in <math>kg/m^3</math>,
<math>P</math> is the total pressure of the system in <math>N/m^2</math>,
<math>\mu_s</math> is the solvent's viscosity in <math>mNs/m^2</math>,
<math>H</math> is the Henry Law constant in <math>1/(Nm^2*(mol fraction))</math>,
and <math>M_s</math> is the molecular weight of the solvent.

A packed towers height can be determined using the equations below when concentration of solute is below 10% so that the assumption that the flow of gas and liquid will be essentially constant throughout the column holds (Towler & Sinnott, 2013). The height of packing <math>Z</math> is given by the following equation:

<math>Z=\frac{L_m}{K_G*a*P}*\int\limits_{y_2}^{y_1} \frac{dy}{y-y_e}\,</math>

where <math>P</math> is the total pressure, <math>a</math> is the interfacial surface area per unit volume, <math>y_1</math> and <math>y_2</math> are the mol fractions of the solute in the gas stream at the bottom and top of the column respectively, <math>G_m</math> is the molar gas flow rate per unit cross-sectional area, and <math>y_e</math> is the mole fraction of solute in the gas that would be in equilibrium with the liquid concentration.

The first half of the equation before the integral can be called the height of an overall gas-phase transfer unit <math>H_G</math> and the second part of the equation is the number of overall gas-phase transfer units or <math>N_G</math>. Using these definitions the above equation can be simplified to

<math>Z=H_G*N_G</math>

These equations assist in sizing an absorption column (Towler & Sinnott, 2013).

==== Film Absorber ====
The final absorber the film absorber is generally used in the case where the heat of absorption must be removed. The film absorber operates by sending the gas and solvent through a heat exchanger where the solvent creates a thin film on the walls of the tubes and the gas flows through the interior allowing for solute transfer. The good heat transfer present in a film absorber makes it preferable for situations where low temperatures are required for a high recovery of the solute (Schmidt 2012).

===Industrial Absorption Processes===
An industrial example is lean oil absorption, which is used to separate nitrogen and other impurities from natural gas. A lean oil is contacted with low quality natural gas, and the methane is selectively absorbed by the lean oil, leaving the impurities behind. The methane is subsequently regenerated from the rich oil as high quality natural gas (Petrogas Systems, 2014).

Other common industrial practices of absorption come from sour gas treatment. Amine gas treating is used to remove hydrogen sulfide or carbon dioxide from gas streams via a reversible chemical absorption. In amine gas treating the sour gas is fed to the bottom an absorber where amine solution is fed to the top along with any necessary make up water. The sour gas components are absorbed into the amine via a chemical absorption method. Sweet gas leaves the top of the absorber whereas the amine out of the bottom, now rich with acidic components is sent to a regenerator where the acid gas components are stripped and the acid gas is generally sent to a flare whereas the amine now lean again is recycled back into the first absorber (Miller & Zawacki, 1978). Figure 5 below shows the typical setup of an amine plant. Another type of sour gas treatment that uses absorption is Merichems LO-CAT process which uses a chelated iron to remove hydrogen sulfide from feed gas in the absorption column (Merichem 2015).
[[File:AmineTreating.png|thumb|400px|center|Figure 5. Amine Gas Treating Plant Schematic]]

==Stripping==
This process separates solutes from solvents (often after absorption, to purify the solvent so that it can be recycled to an absorber). Stripping will depend on the vapor and liquid flow rates, as well as the temperature and pressure of the column. There is a temperature drop down the column, so columns generally have either an increased operating temperature or decreased operating pressure.

The stripping factor of component i is

<math>S_i=K_iV/L</math>

where <math>K_i</math> is the vapor/liquid equilibrium ratio, <math>V</math> is the vapor flow rate entering the column, and <math>L</math> is the liquid flow rate entering the column, will determine how much of solute i will be stripped from the liquid into the vapor phase (Peters & Timmerhaus, 2003). The usual range for the stripping factor is between 1.2 and 2.0, with a stripping factor of 1.4 being most economic.

An example of stripping in industry is the deodorization of food items such as oils. The oil is heated and allowed to trickle down the column while steam flows up from the bottom of the column. At the vapor-liquid interface, volatile components of the oil transfer to the steam and are carried off the top of the column, leaving a purified oil product (Alfa Laval, 2014).

==Bioseparations==
===Importance===
As our ability to manipulate and engineer biological systems improves, biological products are becoming an increasingly important source of therapeutics and fuels. The production of fuels from biomass via either the enzymatic breakdown of a feedstock or the secretion of usable lipids from algae is a promising new energy source. Additionally, enzymes, antibodies and other therapeutic proteins have been applied to the treatment of a wide range of diseases. Although each process requires its own set of separations, all follow the same basic format: separation of biomass, product isolation, and product purification (Belter et al., 1998). This section will provide examples of unit operations in each step. Ultimately, the choice of separation process and unit operations will depend on the specific process and product. The descriptions below are examples of the most common bioseparation operations within the general platform (Harrison et al., 2003).

Bioprocesses begin with fermentations or growth operations. In biofuel production processes, this may involve growing algae or breaking down corn or cellulosic biomass. For the production of therapeutics, mammalian or bacterial cells may be grown in a fermentor and the product secreted into the supernatant or harvested from the cells.

===Biomass Separations===
After fermentation and product production, the solid biomass must first be separated from the desired product. If the product is secreted from the cells, this can be done immediately after fermentation ends. If the product is not secreted, the cells must first be lysed.
Cell lysis is the process of lysing, or breaking, the cell in open. Mechanical lysis is the simplest, and involves physically breaking the cell either by mashing (think mortar and pestle) or blending the cells into a homogenous solution in a homogenizer. Chemical lysis is another method, achieved by introducing an osmotic shock or chemically degrading the cell membrane. Additional separation can be achieved by flocculation, which is the process of aggregating biomaterial by charge neutralization or bridging. These larger complexes are easier to separate from smaller molecules (Harrison et al., 2003).

The next step is removing the unwanted biomass from the product in solution. Separation by centrifugation or sedimentation are the most common, although filtration is sometimes also used for processes where a biomass cake is desired. Both methods utilize density differences to separate the product from the solid biomass (Towler and Sinnott, 2013).

====Sedimentation====

Sedimentation relies purely on the force of gravity, while centrifugation speeds the settling process by subjecting the cells to a centrifugal force. Sedimentation in a settling tank is the simplest method of solid-liquid bioseparation. In this process, biomass in a tank is simply allowed to settle to the bottom over time. While this process is inexpensive, requires little energy and can separate out large volumes of biomass, it generally requires long time periods and is only mostly in very large-scale processes where active centrifugation is difficult (Belter et al., 1998).

====Centrifugation====
Centrifuges are widely utilized across many processes, and thus a wide variety of scales and designs have been developed. Disk-stack centrifuges, in which the solid phase is deposited onto “shelves” in the center of the spinner and liquid phase is pushed to the outside, are some of the most commonly used centrifuges in industry. They are especially suited to biomass separation processes because they can be built on a large scale and are ideal for separating fine solids from liquids. [[File: Disk_stack_centrifuge_towler.png|frame|center|Fig. 6: Diagram of a disk-stack centrifuge (Tolwer et al, 1997).]] Tubular bowl centrifuges are also common and can reach separation efficiencies of up to 90%. Heavier products accumulate along the sides of the bowl, while the light phase flows out the top. They separate products by can be used both to separate solids from liquids and immiscible liquids, such as and oil product and an aqueous broth (Tolwer and Sinnott, 2013). [[File: tubular bowl centrifuge towler.png|frame|center|Fig. 7: Diagram of a tubular bowl centrifuge centrifuge (Tolwer and Sinnott, 2013).]]

Centrifugation scale-up is made easier by sigma analysis, which allows for the estimation of appropriate feed rates for different size centrifuges. The sigma factor is dependent on the inner and outer radius of the centrifuge, the angular velocity, and the sedimentation velocity of the solid particles being separated. It can be thought of as the characteristic cross-sectional area with units of [length]2. The sedimentation velocity can be calculated by

<math>v_g={\frac{2a^2(\rho-\rho_0)}{9\mu}}g</math>

where <math>v_g</math> is the sedimentation velocity, <math>a</math> is the cell or biomass particle diameter, <math>\rho</math> is the particle density, <math>\rho_0</math> is the fluid density, and <math>\mu</math> is the fluid viscosity. The volumetric flow <math>Q</math> can be estimated by

<math>Q=(v_g)(\Sigma)</math>.

The equality

<math>{\frac{\Sigma_1}{\Sigma_2}}={\frac{Q_1}{Q_2}}</math>

can be an easy way to estimate equivalent flow rates between a small-scale centrifuge 1 and larger centrifuge 2 (Harrison et al., 2003).

====Example: Centrifugation Scale-up====

You are trying to separate a cell of radius 0.4 <math>\mu</math>m with a density of 1.05 g/cm3 from broth of mostly water (density of 1 g/cm3 and viscosity of 0.01 g/cm s). The sigma factor of the centrifuge you are using is 1 x 106 cm2. A] What volumetric flow rate should you use? B] If you want to scale up the process to a centrifuge with <math>\Sigma</math> = 3 x 106 cm2, what flow rate would you use in the larger centrifuge?

Solution:
A] Using the equation for <math>v_g</math>, and being mindful of units, the sedimentation velocity equals 1.74 x 106 cm/s. The flow rate, then, equals

<math>Q=(1.74 x 10^-6)(1,000,000) = 1.74 cm^3/s = 0.104 L/min</math>.

B] Keeping in mind that for the same process, <math>v_g1 = v_g2,</math> and rearranging the sigma factor equality, the new flow rate is

<math>Q_2 = {\frac{\Sigma_2 x Q_1}{\Sigma_1}} = {\frac{(3 x 10^6)(0.104)}{1 x 10^6}} = 0.313 L/min </math>

===Product Isolation===
Liquid-liquid separation, to extract the product from the aqueous phase, is much less straightforward than liquid-solid extraction. Many methods - especially adsorption, filtration, and precipitation - are similar in principle to operations found in other, non-biological separations. The exact separations used depend on the nature of the product and the scale of the process. These processes are nearly identical to their non-biological counterparts, and their description is left to other sections.

Particular care needs to be taken with protein products because of their instability, and the selection of an appropriate solvent or adsorbent is crucial to a successful process (Harrison et al., 2003).

===Product Purification===
The final steps of protein purification and polishing remove any remaining contaminants and bring the concentration of product to an appropriate value for applications. Purification processes for food-grade and medical products can be extensive, as sterility and high purity are essential. Purification in fuel-producing processes may be less extensive, depending on the process. Chromatography and crystallization are two common steps in purification and are especially used in industrial scale protein production. Several different types of chromatography exist with the ability to carry out different types of separations.
Chromatography is similar to adsorption in that it relies on differences in affinity between solutes and a solid surface. A solution is eluted through a column containing a solid resin with various affinities for the substances in solution. In adsorption, the solutes are evenly saturated throughout the column. Chromatography differs in that solutes are deposited a resin phase before the column is flushed with an elution solvent specific that results in solutes eluted in bands.

==== Ion Exchange Chromatography ====

There are two main types of ion exchange columns—anion and cation. Anion exchange resins have a positive charge and are used to retain products with a negative charge. Cation exchange resins have a negative charge and are used to retain products with a positive charge. The pH of the elution buffer is change to force a specific solute to wash out, depending on whether the pH of the buffer is above or below the isoelectric point of the solute (Belter et al., 1998). This is especially useful for the separation of protein product (including antibodies), nucleic acids, and other charged molecules. When the solutes have sufficiently different isoelectric points, the pH of the buffer is manipulated to affect the solute charge and force the product to elute while the solute remains preferentially bound to the resin, or vice versa (Harrison et al., 2003). In general, the most strongly charged molecules will remain in the column for a longer period of time. Elution washes through the weakly bound ions before the more strongly bound ions. Different speeds of elution can be visualized as in figure 8.

[[File:chromatography.png|frame|center|Fig. 8: Illustration of product bands in an elution chromatography column (Belter et al., 1998).]]

==== Size Exclusion Chromatography ====

In gel filtration chromatography, small molecules are "trapped' by the porous resin and take longer to flow through the column. Larger products will elute first because the smaller molecules are better able to penetrate the resin. This forces them to take a much longer path through the column, which means it takes longer for them to elute. This operation is often used when there is a distinct difference in size between the desired product and other solutes.

==== Affinity Separations ====

Affinity chromatography is very similar to ion exchange chromatography in that the interactions between the material in the column and the molecules in the feed. The main difference is that affinity chromatography can rely on a great variety of types of interactions. Two very common types of affinity are exploited in affinity chromatography columns. The first is immunoaffinity. Proteins are specifically bound by antibodies which can be incorporated onto beads and used in chromatography. Antibodies are designed to bind only a single protein, so these interactions are considered to be highly specific. The protein can be eluted using a buffer that changes the pH or salinity in the column, which adversely affects binding (Hage, 1999).

Proteins can be separated from very complex and unrefined mixtures using this method because of the great specificity. Antibody-protein interactions are so specific that individual proteins can be isolated from blood samples as shown in the figure below. The main drawback of this method is simply that specific interactions between proteins and binding targets do not always exist. Antibodies exist for many proteins, but for others they must be customized or created using some sort of evolution process. This is not a trivial task and can require a significant capital investment (Hage, 1999).

[[File:Affinity_Chromatography_Example.PNG|frame|center|Fig. 9: Affinity purification separating fibrinogen from human plasma using an anti-fibrinogen antibody (Hage, 1999).]]

The other main type of affinity chromatography is based on protein specific tags and the molecules or surfaces to which they bind. One of the most common types of protein tags used is the polyhistidine tag. This tag consists of 6-8 consecutive histidine residues which can be added to the exterior of the desired protein product. The addition of this tag requires alterations to the coding sequence of the protein. The polyhistidine tag binds strongly with nickel and cobalt ions. The product with the tag can then be eluted with imidazole—a small molecule with the same structure as the functional group of the amino acid histidine. Imidazole will bind the cobalt and nickel ions more strongly than the histidine in the tag. Along with chromatography, protein tag interactions can be leveraged with the use of beads that can be deposited directly into the solution containing the protein of interest (Lichty et al., 2005).

Different proteins can be purified in this manner with varying levels of efficiency. There is a very high dependence on the size of the protein, electronic properties, and steric considerations. The figure below demonstrates varying separation efficiencies with proteins of different sizes (HisPur Cobalt Resin - Thermo Fisher Scientific, 2016).

[[File:HisTagPurification.PNG|frame|center|Fig. 10: Protein gel demonstrating the separation of proteins from a sample. The far left column represents the most efficient separating conditions using cobalt and the far right column represents a negative control with no purification performed (HisPur Cobalt Resin - Thermo Fisher Scientific, 2016).]]

Several other types of tag-bead interactions can be utilized in separations processes. Maltose Binding Protein is a small protein that can be added to a protein of interest. It binds strongly with beads coated in immobilized maltose and can be released by flushing with maltose. As MBP is a full sized protein that typically must be removed from the protein of interest in order for it to be used. In this case, the site specific TEV protease is often used cleave MBP from the protein of interest. In addition, under specific circumstances, other unique tags can be used and provide varying levels of specificity in separations. The Flag tag, 3x Flag tag, Glut tag, and Strep tag. While these are all commonly used, the polyhistidine tag is the most popular because it gives the highest level of specificity. Some of the main purification tags are compared in the table below (Lichty et al., 2005).

{| class="wikitable"
|-
! Tag
! Size (aa)
! Resin
! Eluting Agent
! Capacity (mg/ml)
! Cost/10 mg
|-
| His
| 6
| Talon
| Imidazole
| 5-14
| $18
|-
| Maltose Binding Protein
| 396
| Amylose
| Maltose
| 3
| $12
|-
| Glutathione S-Transferase
| 218
| GSH-Sepharose
| Glutathione
| 10
| $36
|-
| Strep II
| 8
| Strep-Tactin Sepharose
| Desthiobiotin
| 0.1
| $293
|-
| FLAG
| 8
| Anti-FLAG M2 MAb Agarose
| Flag peptide
| 0.6
| $1045
|}

Unfortunately, based on the contents of the solution, the characteristics of the protein, and the specific product being used, there are extremely large amounts of variation in yield and purity for each type of tag. Therefore, it is crucial to allocate time towards preliminary tests and scale-up experiments with the exact solution from which the protein of interest is being separated. The his tag is typically the most used because, as seen in the table above, it has one of the lowest costs and highest yields. In addition, because it is only a 6 amino acids, it often is not cleaved from the protein of interest which minimizes the number of steps required in the overall purification process (Lichty et al., 2005).

==== Crystallization ====

Crystallization, or the formation of solute crystals from a solution, is especially useful in biomolecule separations because it is possible to obtain a 99.9%+ product purity. In crystallization, a diluent is added to the homogeneous solution that reduces the solubility of the product to the point that it “falls out” of solution and crystallizes. It is similar to precipitation but results in the formation of crystals rather than unordered aggregates. Crystallization can be used on a laboratory scale for determining protein structure, on on the industrial scale for antibody and therapeutic protein productions. Batch crystallizers are often used in industry because of their simplicity and inexpensiveness compared to continuous crystallization (Harrison et al., 2003).

==Membrane Separation==
Membrane separation takes advantage of the selective permeability of membranes; they allow certain particles to pass through and selectively stop other, generally unwanted, particles. The component that passes through is called the permeate and the component stream that is rejected is called the retentate or concentrate. The applicability of membranes comes from the fact that their selectivity is determined by their pore size, which can be controlled during the creation of the membranes. Additionally, Membrane processes do not require heat meaning they generally require less energy than conventional separations technology such as distillation and crystallization. Membrane separations processes are generally classified as microfiltration, ultrafiltration, or nanofiltration depending on the size of the particles to be filtered out.

[[File:membrane techs.png|frame|Figure. Cutoffs for different membrane categories]]

===Membrane Selection, Construction, and Flow Geometries===
Membrane permeability and selectivity are the two most important factors to consider when selecting a membrane. For gas separations, the permeation of the gas is usually facilitated by the gas dissolving in the membrane on one side and then evaporating on the permeate side. Therefore permeability depend largely on the solubility of components in the membrane.

The two most commonly utilized membrane configurations are hollow fiber and spiral wound. Hollow fiber is generally the most commonly utilized module for gas separations. These are formed by gluing the two ends of the hollow fiber to a resin forming a closure. The fibers are housed in a shell much like a heat exchanger. The feed flows past thousands of tubes with the permeate flowing into the hollow tubes and out the closure. The retentate then flows out of the shell not having gotten through the membrane.

[[File:hollow membrane.jpg|frame|Figure. Hollow fiber membrane module]]

Spiral wound membranes are created by sealing two membrane sheets back to back on three edges to form a sort of pocket. This fourth open edge is then attached to a porous tube which allows permeate to go through it. Several membrane pockets are attached to a single tube and wrapped around in a spiral.

[[File:spiral membrane.jpg|frame|Figure. Spiral wound membrane module]]

Flow geometry is usually either dead-end geometry or cross flow geometry. In dead end, the fluid flow is normal to the membrane surface while cross flow is parallel to the membrane surface. Dead end geometry is usually used with hollow fiber membranes while cross flow is used with spiral wound membranes. Each geometry has advantages and disadvantages. Dead end geometry is generally cheaper to set up and therefore has lower initial capital costs. However, it is very vulnerable to membrane fouling, which reduces the effectiveness of the membrane. This is usually the geometry set up for small scale lab experiments. The tangential flow devices are more cost and labor-intensive, but they are less susceptible to fouling due to the sweeping effects and high shear rates of the passing flow. Most commercial industrial membrane separations are done using spiral wound cross flow membrane modules.

===Applications===
====Food Industry====
Due to the fact that MD can be conducted at relatively low feed temperatures, it was successfully tested in many areas where high temperature applications lead to degradation of the process fluids especially in food processing. It was demonstrated that MD can be used for the concentration of milk, for the recovery of volatile aroma compounds from black currant juice, and for the concentration of many other types of juices including orange juice, mandarin juice, apple juice, sugarcane juice, etc.
====Reverse Osmosis====
Reverse osmosis is the most widely used membrane separation process. In this process, fresh water passes through the membrane while dissolved salts and other solids are rejected and stay in the concentrate. In this process, feed water is pressurized in order to overcome the osmotic potential difference between the salty retentate and the fresh water desired. These processes are generally run using spiral wound membrane cylinders using a cross flow setup.

===Membrane Model===
The two most important components when considering different membranes are the permeability, which will determine flux through the membrane, and selectivity, which will determine what passes through the membrane and how much. The flux through a membrane is defined as:
<math> M_i = \frac{P_i}{δ}(p_{i,f} - p_{i,p})</math>
Where Mi is the molar flux of component i, Pi is the permeability of the membrane for component i, δ is the membrane thickness, and pi,f and pi,p are the partial pressures of component i on the feed side and permeate side respectively.
The average flux across a long cylindrical membrane such as the spiral wound module is given by:
<math> \int_0^Lm \frac{M_i,dx}{L_m}</math>
Where Lm is the length of the cylinder and x is length in meters

Membrane selectivity of the ideal separation factor is given as the ratio of the permeability of one substance over another as shown:

<math> S_(i,j) = P_i/P_j </math>

where Sij is the selectivity of the membrane for component i over j.

==Cyclones==
Centrifugal Separators, more commonly know as Cyclones, are one of the most widely used gas-solid separators. They are typically used for particles that are 5μm or larger, but if agglomeration occurs they can sometimes separate particles as small as particles as small as .5μm. These cyclones can be designed for high efficiency separations or for high throughput (Towler, 2012).

[[File:cyclone.png|thumb|center|Reverse-flow Cyclone (Towler, 2008)|500x300px]]

===Design===
Design Procedure
There are multiple design methods for cyclones. The deign method discussed below is the Stairmand method. This involved creating two designs one of which was for high throughput and the other which was designed for high efficiency. Using these base designs future cyclones can simply be estimated using scaling factors (Towler, 2012).

1. Select whether efficiency or performance is more important for your design

2. Determine the particle size distribution in your stream

3. Estimate the number of cyclones needed

3. Calculate the cyclone diameter

4. Calculate the scale up factor

5. Calculate the cyclone performance and overall efficiency

6. Calculate the cyclone pressure drop

7. Cost the system

====Particle Size====
The diameter of particles separated by cyclones are governed by this design equation. This equation is dependent upon the standards of the cyclone and how much you are trying to deviate from its standard operating conditions. Looking into the equation it shows that a larger internal cyclone diameter, higher flowrate, higher difference in density, and lower viscosity result in smaller particles being seperated. The whole term after d1 is known as a scale up factor. When designing a cyclone you would determine the flow rate, viscosity, particle size, and density difference would be known prior to using this design equation, so this would be used to optimize the diameter of the cyclone you were designing (Towler, 2012).

[[File:Towler 10.8.PNG|center|Reverse-flow Cyclone (Towler, 2012)|500x300px]]

d1=mean diameter of particle separated at the standard conditions, at the chosen
separating efficiency

d2= mean diameter of the particle separated in the proposed design, at the same
separating efficiency

Dc1 =diameter of the standard cyclone

Dc2 =diameter of proposed cyclone

Q1 =standard flow rate

Q2 =proposed flow rate

Δρ1 =solid-fluid density difference in standard conditions

Δρ2 =density difference, proposed design

μ 1=test fluid viscosity

μ 2=viscosity, proposed fluid.

[[File:Towler10.44.PNG||thumb|center|(a) Standard Cyclone Dimensions (b) High Gas Rate Cyclone (Towler, 2012)|500x300px]]

====Efficiency====
Efficiency is the percentage of total solids that are removed from the stream The graphs below were experimentally for the two standard designs Stairmand Method. This graph has the efficiency vs the particle size (Towler, 2012).

[[File:Towler10.45.PNG||thumb|center|Performance curves, standard conditions. (a) High-efficiency cyclone performance curves, standard conditions. (b) High gas-rate cyclone (Towler, 2008)|600x400px]]

If you have to scale the efficiency it can be estimated by the point where a horizontal line from the efficiency of the original particle size intersects the new particle size. A demonstration is on the graph below.

[[File:Towler 10.46.PNG||thumb|center|Scaled Performance Curve(Towler, 2008)|500x300px]]

====Pressure Drop====
One of the most important design factors in a cyclone is its pressure drop. The pressure drop occurs because entry and exit losses, kinetic energy losses, and friction losses that occur inside the cyclone. The pressure drop can be calculated using the following equation (Towler, 2012).

[[File:Towler 10.9.PNG|center|Reverse-flow Cyclone (Towler, 2008)|500x300px]]

The variables are defined as follows:
DP = cyclone pressure drop, millibars;
rf = gas density, kg/m3
;
u1 = inlet duct velocity

u2 = exit duct velocity

rt = radius of circle to which the center line of the inlet is tangential

re = radius of exit pipe

φ =factor from figure below

ψ = fc(As/A1)

fc = friction factor

As = surface area of cyclone exposed to the spinning fluid

A1 = area of inlet duct

[[File:Towler_10.47.PNG|thumb|center|Cyclone Pressure Drop Factor (Towler, 2008)|600x400px]]

====Cost Estimation====
The following values are represented in 2002 dollars. Cost is mostly a function of the amount of gas that need to be passed through the cyclone and the units are in standard meters cubed per second (sm3/sec).

The costs can be broken down into three categories:

Capital Costs: $4600-$7600 sm3/sec

Operations and Maintenance: $1500-$18,000 sm3/sec

Annualized Cost: $2800-$29000 sm3/sec

This all comes to $.47-$440 per metric ton of pollutant removed. The high variation in these costs have to do with smaller units being much less efficient per unit of pollutant removed. Furthermore the cost is also effected by the amount of particulates that are in the gas to begin with. For example a gas with 10 g/sm3 and air with 100sm g/sm3 would have similar pumping costs, but very different cost per unit of solids removed (EPA).

==Other Separation Processes==
===Extraction===
Liquid-liquid extraction is a process for components with overlapping boiling points and azeotropes. The process requires a solvent such that some of the components of the mixture are soluble, and then the components will be separated based on this solubility in the liquid. This process can operate at moderate temperatures and pressures, so is not very energy intensive. However, a distillation column is required to extract the solvent for recycle. More recently, supercritical fluids have replaced liquid solvents in some processes for L/L extraction, due to the solute’s ability to more rapidly diffuse through them. The issue with these fluids, however, is that they must be operated at extremely high pressures and temperatures, increasing both capital and operating expenses of the process (Peters & Timmerhaus, 2003).

===Crystallization===
This process recovers solutes that have been dissolved in solution. The resulting product is in the solid phase. Depending on the material properties of the solute and solvent, the solute is recovered by precipitation after cooling, removal of solvent, or adding precipitating agents. Crystallizers are designed based on phase equilibria, solubilities, rates and amounts of nuclei generated, and rates of crystal growth. Every crystallization process is a unique system, so plant evaluation is usually required before complete implementation. Crystallization can be performed in both batch and continuous processes, and design features can control crystal size to an extent (Peters & Timmerhaus, 2003).

===Membrane Separation===
This separations process uses selectively permeable membranes to separate components in a mixture. Typically, one of the components will freely pass through the barrier while the other components will not. The stream that passes through the membrane is the permeate and the stream that does not pass is the retentate. The driving force behind this separation is a pressure gradient. Membrane separation is beneficial because it can separate mixtures at the molecular and small particle level. Furthermore, there is no phase change required so the energy input is low. Limitations of this process include achieving high product purity, incompatibility with certain stream components, low operating temperature, and low flow rates. Although membrane separation is generally not scaled up, examples of scaled-up membrane separation include seawater desalination and hydrogen recovery (Peters & Timmerhaus, 2003).

===Adsorption===
Adsorption involves an adsorbent and adsorbate. The adsorbent is typically a solid, and will typically separate the adsorbate from the stream. This process usually includes a desorption step that regenerates the adsorbent for further use. Raising the temperature or increasing the concentration of the adsorbate can reverse the adsorption process. Although the recycle of the adsorbent is a very economic design feature, the downside of this step is that it results in a cyclic process, which introduces complexity to the overall process. Industrial applications of this process are for bulk separations and gas purification. The adsorption/desorption process in these situations involves a large amount of heat transfer, which design engineers must take into account when sizing and selecting equipment material (Peters & Timmerhaus, 2003).

===External Field/Gradient Separation===
These separations use external force fields or temperature gradients to separate responsive molecules or ions. The use of these processes is fairly limited to a few specialized industrial applications (Peters & Timmerhaus, 2003).

===Settling and Sedimentation===
In settling processes, solid particles or liquid drops are separated from a stream by gravity. The stream can be in either the liquid or gas phase. For vapor-liquid mixtures, flash drums are generally used to separate the mixture. The velocity of the vapor must be less than the settling velocity of the liquid drops for this separation to occur. For liquid-liquid separation, the horizontal velocity of the fluid must be low enough to allow the low-density droplets to rise to the interface and the high-density droplets to move away from the interface and coalesce. In sedimentation, the result of the process is a more concentrated slurry. Typically a flocculating agent is used to aid in the settling process. One way to perform this separation is to use a cone-shaped tank with a slowly revolving rake that scrapes and moves the thickened slurry to the center of the cone for removal (Peters & Timmerhaus, 2003).

[[File:Example.jpg]]

====Clarifiers====
[[File:Circular_Clarifier.png|300px|thumb|bottom|Figure 9: Circular clarifier with some components labelled.]] [[File:Rectangular_Clarifier.png|300px|thumb|bottom|Figure 10: Rectangular clarifier with some components labelled.]]

Clarifiers are one of the methods used for the continuous removal of particulate solids from liquids through sedimentation by gravity. Applications include process water pretreatment, waste water treatment, and drinking water purification. Historically, clarifiers were originally developed to limit nutrient input into surface water due to fear of eutrophication. Today, they have a number of uses, particularly in wastewater treatment processes, metal removal, disinfection, and membrane pretreatment. The process helps removed dissolved solids, silt, and undesirable metals from the water, making it more suitable for downstream processes as well as human consumption (Wilson, 2005).

Clarifiers are typically used in conjunction with coagulation or flocculation agents, which promote dissolved particles to join into clumps and settle out of solution (Towler and Sinnot, 2012). Clarifiers typically consist of a large circular tank with a rotating rake at the base which scrapes settled solids towards the center. In the case of a rectangular clarifier, they are scraped to one side. Diagrams of both are represented in figures 9 and 10, respectively (NMED Surface Water Quality Bureau, 2015). Separated solids are allowed to settle to the bottom of the tank as a sludge, whereupon they are collected by the rake and disposed of properly. In the case of floating contaminants, it is possible for the clarifier to include a skimmer as well.

Clarifier efficiency varies with certain factors, including the settling characteristics of solids removed and the surface overflow rate of the tank. Clarifier efficiency can be found using the following relation:

<math>\begin{align}
E_{TSS} &= E_{TSSmax}\left ( 1 - e^\frac{\lambda}{SOR} \right )
\end{align}</math>

where <math>E_{TSS}</math> is the efficiency of total suspended solids (TSS) removal, <math>E_{TSSmax}</math> is the maximum possible efficiency, <math>\lambda \left [\frac{m}{d} \right ]</math> is the settling constant, and <math>SOR \left [\frac{m^3}{m^2 d} \right ]</math> is the surface overflow rate. The effect of flocculation chemicals on TSS can be seen in figure 11. However, it should be noted that chemical addition will increase sludge quantity and may have an adverse effect on plant aesthetics, which increases maintenance costs (Wilson, 2005).

[[File:Chem_Addition.png|200px|thumb|bottom|Figure 11: The effect of flocculating agents on total suspended solids removal in clarifiers.]]

=====Lamella Clarifiers=====

Lamella clarifiers use inclined plates in order to maximize the settling area for solids. Solids continue to settle into a hopper at the bottom of the tank while clarified water exits up through the inclined plates. This allows for the design of a smaller tank, which leads to large savings in capital costs. A lamella clarifier is pictured in figure 12.

[[File:Lamella_Clarifier.png|300px|thumb|bottom|Figure 12: A lamella clarifier with components labeled.]]

Typically, inclined plates are installed at an angle of 45 to 60 degrees and spaced 40 to 120 mm apart, which increases effective settling surface area by a factor of 6 to 12 compared to traditional clarifiers. For effective use, it is recommended that the Reynolds number be below 2000, Froude number higher than 10-5,and detention time be longer than 3 to 5 minutes. For this implementation, the equations are as follows:

<math>\begin{align}
N_{Re} &= \frac{VR}{\nu} \\
N_{Fr} &= \frac{V^2}{Rg}
\end{align}</math>

where <math>R</math> refers to the hydraulic radius, which is the cross-sectional area of the lamella, <math>V</math> is the liquid velocity, <math>\nu</math> is the kinematic viscosity, and <math>g</math> is the gravitational constant (Wilson, 2005).

=====Advantages=====

Clarifiers offer a proven, relatively inexpensive solution for solids removal. The chemical coagulants used are cheap and provide a low operating cost as well as simple maintenance. Construction is typically simple, leading to low capital costs and equipment that is easy to accommodate and maintain. Their design is also flexible, with various options such as skimmers and scrapers offering increased removal efficiency (Wilson, 2005). Operation of clarifier tanks also has lower energy requirements than membrane filtration for solids removal, given that most of the separation is aided by gravity. Water exiting clarifier units has a silt density index (SDI) averaging 4.0, which is low enough for further membrane treatment such as reverse osmosis (Prihasto, 2009).

=====Disadvantages=====

Clarifiers necessitate low turbulence to prevent resuspension of solids. This essentially requires a low entrance velocity, which can limit the production rate of certain processes or call for more clarifier units, which would drive up costs. Furthermore, clarifiers require frequent cleaning before sludge becomes too difficult to remove and reduces effectiveness. In the case of lamella clarifiers, sludge buildup on the inclined plates results in uneven flow distribution which could harm efficiency (US EPA, 2003). For this reason, maintenance requirements for lamella clarifiers are higher, but they can be reduced through the implementation of removable plates (Wilson, 2005). Clarifiers also only remove solids, so pH will not be affected, leading to the need for further pH adjustment (NMED Surface Water Quality Bureau, 2015).

=====Clarifier Design Calculations and Typical Design Values=====

======Detention Time======

Detention time (DT) is the time is takes for a unit of water to travel from the inlet of the clarifier unit to the outlet. During typical operations, the design value for this is 2 to 3 hours.

<math>\begin{align}
DT &= \frac{Tank\ Volume}{Influent\ Rate}
\end{align}</math>

======Surface Overflow Rate======

Surface overflow rate (SOR) measures the flow into the clarifier per square foot of surface area. Typical design values are 400 to 800 gal/day/sq. ft.

<math>\begin{align}
SOR &= \frac{Volumetric\ Flow\ Rate}{Surface\ Area}
\end{align}</math>

======Weir Overflow Rate======

Weir overflow rate (WOR) describes the flow in gallons per day per linear foot of weir. Typical values are 10,000 gal/day/ft.

<math>\begin{align}
WOR &= \frac{Volumetric\ Flow\ Rate}{Weir\ Length}
\end{align}</math>

======Solids Loading Rate======

Solids loading rate (SLR) describes the mass of solids in the clarifier influent per square foot of surface area. This value should not exceed 30 lbs/day/sq. ft.

<math>\begin{align}
SLR &= \frac{Solids\ Mass\ Flow\ Rate}{Surface\ Area}
\end{align}</math>

===Flotation===
Flotation is a process designed for specific solid-solid mixtures. It works by generating gas bubbles in a liquid that attach to selected solid particle. Afterwards, the particles rise to the liquid surface where they are removed by an overflow weir or mechanical scraper. The separation depends on the surface properties of the particles and its preference to attach to the gas bubbles. To meet the necessary requirements of the flotation process, a number of additives can be used to control things like the pH of the liquid-solid mixture, the activity of the solid surface, and the froth that can assist in separation. The bubbles can be produced by gaseous dispersion, dissolution, or electrolysis of the liquid (Peters & Timmerhaus, 2003).

===Centrifugation===
This process is similar to external field separation in that an external force field is applied to separate a mixture. When gravity separation is too slow due to particle densities, particle size, settling velocity, or the formation of an emulsion, centrifugation is commonly used. Centrifugal force increases the total force acting on the particle and results in faster separation times. This process is generally used to separate solids from liquids, however it can also be used to separate two liquids with very different densities (Peters & Timmerhaus, 2003).

===Drying===
Drying is performed to remove liquid from a liquid-solid mixture and produce a dry solid. Water is most often the liquid removed, but organic liquids are removed from solids on occasion as well. The heat required to vaporize the liquid is usually obtained by a series of gas-solid contacting devices. Feed condition and temperature sensitivity of the solid dictate the type of contacting device that is used. There are two groups of dryers that differ by the dependence of either mechanical means or fluid motion for gas solid contact. Another feature of dryers is to use either direct (hot gas) or indirect (conductive surface) heating (Peters & Timmerhaus, 2003).

===Evaporation===
Evaporators separate solvents from a solution by evaporation. The difference between evaporation and distillation is that evaporation requires the solute be nonvolatile. Because of this, a high separation can be achieved with one stage. Evaporators are essentially reboilers, so evaporation is a very energy-intensive process with a high thermal economy (Peters & Timmerhaus, 2003).

===Filtration===
Filtration is a process that separates a mixture of solid in a liquid or gas by passing the mixture through a porous medium in which the particles do not pass. Filtration is done by either cake filtration (particles found on the surface of the filter) or depth filtration (particles found within the filter). Cake filtration is generally performed with a cloth as the filtration medium (Peters & Timmerhaus, 2003).

==Conclusion==
Separation is a key part of most chemical processes, and there is a great variety of techniques to perform separation of compounds based on size, volatility, charge, and many other features. A common technique with which the process engineer should be familiar is distillation, but he or she should also be aware of the other available options. Some techniques may be less expensive, less energy-intensive, or more effective than distillation, depending on the specific separation problem. Therefore, the separation strategy should be carefully considered.

==References==
Air Control Technology Fact Sheet. Enviornmental Technology Fact Sheet. http://www3.epa.gov/ttncatc1/dir1/fcyclon.pdf

Belter PA, Cussler EL, Hu WS. Bioseparations: Downstream Processing for BIotechnology. New York: John Wiley; 1998.

Biegler LT, Grossmann IE, Westerberg AW. Systematic Methods of Chemical Process Design. Upper Saddle River: Prentice Hall; 1997.

Danckwerts P (1965) The Absorption of Gases in Liquids. Pure and Applied Chemistry UK 10:625-642.

Development Document for the Final Effluent Limitations Guidelines and Standards for the Metal Products and Machinery Point Source Category (Report). US Environmental Protection Agency. 2003.

Erwin, D. Industrial Chemical Process Design. New York: McGraw Hill, Professional Engineering; 2002.

Hage, David S. 1999Affinity Chromatography: A Review of Clinical Applications. Clinical Chemistry 45(5): 593–615.

Harrison RG, Todd P, Rudge SR, Petrides, DP. Bioseparations Science and Engineering. New York: Oxford University Press; 2003.

Hage, David S. 1999Affinity Chromatography: A Review of Clinical Applications. Clinical Chemistry 45(5): 593–615.

HisPur Cobalt Resin - Thermo Fisher Scientific N.d. https://www.thermofisher.com/order/catalog/product/89964, accessed February 21, 2016.

Lamella Plate Clarifier. Hydro International Web site. Available at: http://www.hydro-int.com/uk/products/lamella-plate-clarifier?s=0&r=uk. Accessed February 2, 2016.

Lean Oil Absorption. PetroGas Systems Web site. Available at: http://petrogassystems.com/technology/natural-gas-processing-and-dew-point-control/lean-oil-absorption. Accessed February 19, 2014.

Lichty, Jordan J., Joshua L. Malecki, Heather D. Agnew, Daniel J. Michelson-Horowitz, and Song Tan 2005Comparison of Affinity Tags for Protein Purification. Protein Expression and Purification 41(1): 98–105.

Merichem Gas Technologies. ®LO-CAT PROCESS available at http://www.merichem.com/images/casestudies/Desulfurization.pdf Accessed 6 Feb. 2015.

Miller L.N. & Zawacki T.S. , US 4080424, "Process for acid gas removal from gaseous mixtures", issued 21 Mar 1978, assigned to Institute of Gas Technology

NMED Surface Water Quality Bureau, New Mexico Water Systems Operator Certification Study Manual, New Mexico Environment Department, 2015.

Peters MS, Timmerhaus KD. Plant Design and Economics for Chemical Engineers. 5th ed. New York: McGraw Hill; 2003.

Prihasto, N; Lui, Q; Kim, S. Pre-treatment strategies for seawater desalination by reverse osmosis system. 2009; 249(1): 308-316. doi:10.1016/j.desal.2008.09.010

Schmidt Eberhard (2012) Waste Gases, Separation and Purification. Ullman’s Encyclopedia of Industrial Chemistry Germany 2:174-181.

Seider, W.D., Seader, J.D., and Lewin, D.R. (2004). ''Process Design Principles: Synthesis, Analysis, and Evaluation.'' New York: Wiley.

Stripping Column. Alfa Laval Web site. Available at: http://www.alfalaval.com/solution-finder/products/soft-column/Documents/Stripping%20Column.pdf. Accessed February 19, 2014.

Towler G, Sinnott R. Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013.

Turton, R.T., Bailie, R.C., Whiting, W.B., and Shaewitz, J.A. (2003). ''Analysis, Synthesis, and Design of Chemical Processes'' Upper Saddle River: Prentice-Hall.

Wankat, P.C. (2012). ''Separation Process Engineering.'' Upper Saddle River: Prentice-Hall.

Wilson, T.E., Clarifier Design, 2nd Ed., McGraw-Hill: New York, 2005.

Separation processes

2016-02-22T01:33:02Z

Bts511: /* Cyclones */

Authors: Nick Pinkerton, [2014] Karen Schmidt, [2014] James Xamplas, [2014] Emm Fulk, [2015] and Erik Zuehlke, [2015] John Dombrowski [2016] , Brett Sleyster [2016] , Robert Cignoni [2016] , and Osman Jamil [2016] 

Stewards: David Chen, Jian Gong, and Fengqi You

Date Presented: February 9, 2014 /Date Revised: February 1, 2014

 

==Introduction==
Essentially all chemical processes require the presence of a separation stage. Most chemical plants comprise of a reactor surrounded by many separators. Separators have a countless number of jobs inside of a chemical plant. A separator can process raw materials prior to the reaction, remove incondensable gases, remove undesired side products, purify a product stream, recycle materials back into the process, and many other jobs that are essential to the process.

Chemical engineers must understand the science of separation and the variety of ways that separation can take place. There are many ways to perform a separation some of these including: distillation, absorption, stripping, and extraction. The science of separation revolves around the presence of two phases that are in contact and equilibrium (Wankat, 2012).

[[File:Sepmeth.JPG|frame|Figure 1. Separation methods by property]]

==Theory==
===Vapor-Liquid Equilibrium===
Separation processes are based on the theory of vapor-liquid equilibrium. This theory states that streams leaving a stage in a separation process are in equilibrium with one another. The idea of equilibrium revolves around the idea that when there is vapor and liquid in contact with one another they are in constantly vaporizing and condensing. Different components in the mixture will condense and vaporize at different rates. There are three types of equilibrium conditions that can be subdivided into thermal, mechanical and chemical potential categories. These separate equilibrium states are given as:

<math>T_{liquid} = T_{vapor}</math>

<math>p_{liquid} = p_{vapor}</math>

<math>chemical potential_{liquid} = chemical potential_{vapor}</math>

==Distillation==
===Flash Distillation===
Flash Distillation is one of the simpler separation processes to be employed in a chemical plant. The main premise of flash distillation is that a portion of a liquid feed stream vaporizes in a flash chamber or a vapor feed condenses. Vapor-liquid equilibrium will cause the vapor phase and the liquid phase to have different compositions. The more volatile component of the mixture will compose of a larger portion of the vapor. This simple separation is easy to manufacture but does not result in large degrees of separation.

Flash distillation requires a feed stream that is pressurized and heated and then passed through a valve into a flash drum. The large pressure drop across the valve will result in a partial vaporization of the fluid. Vapor will be removed overhead from the flash drum while the remaining liquid will collect at the bottom of the drum and be removed. Most flash drums will contain an entrainment eliminator which is a screen that prevents liquid from being carried into the vapor effluent. Figure 2 shows a simple overview of the flash distillation process. As shown, there is a heater that flows into a let-down valve where the two-phase flow begins. Variables y and x are the mole fractions of the more volatile component in the vapor and liquid effluents, respectively.

[[File:Flash.gif|center|frame|Figure 2. Flash Distillation Flow Diagram]]

===Column Distillation===
Distillation columns are the most widely used separation technique used in the chemical industry, accounting for approximately 90% of all separations (Wankat, 2012). Distillations in columns consist of multiple trays that each act at their own equilibrium conditions. Large columns are able to perform complete separations of binary mixtures as well as more complex multi-component mixtures.

[[File:column.jpg|250px|center|]]
===Stages===
Columns are separated into stages by the presence of trays. These trays allow for vapor-liquid contact and equilibrium to occur. Typically, the more stages in a column, the larger separation that can be achieved. There are many different types of trays that can be used in a column.
====Sieve Trays====
The simplest and least expensive tray type is the sieve tray which is a sheet of metal with holes punched into it to allow vapor flow. Sieve trays can have different hole patterns and sizes that will affect the tray efficiency and flow rates.

[[File:sieve.jpg|200px|center|]]

====Sieve Tray Design Procedure====

The design of these plates is done through a trial-and-error process. Most commercial process simulations (such as HYSYS) have default tray designs, and automatically specify dimensions. However, these dimensions selected or calculated by the simulations may not give the best performance for your system, so it is valuable to understand how to design the sieve trays and how specific parameters may affect performance. Hand calculations using the following methods can be used to guide the simulation programs to better design. This section will use sample data to work through an example of the process. The following is a general list of steps for designing a sieve plate:

=====1. Calculate the maximum and minimum vapor and liquid flow rates for the turndown ratio required.=====
This data can be collected from a McCabe-Thiele diagram and/or from process simulation data.

Data from McCabe Thiele diagram, for example: 
Number of stages = 10 
Slope of top operating line = 0.185 
Slope of bottom operating line = 1.43 
Top composition = 98.8 mol% acetone 
Bottom composition = 4 mol% acetone (a) 
Minimum reflux ratio = 0.31 

Collect data rom mass balances: 
MW of feed = (0.6 kmol a/kmol)*(58 kg/kmol a) + (0.4 kmol acetic acid/kmol)*(60 kg/kmol aa) = 59 kg/kmol 
Feed rate = 100 kmol/h 
Using material balances... D = 57.3 kmol/h ; B = 42.7 kmol/h 
Vapor rate, V = D(1+R) = 57.3 kmol/h (1+.31) = 75.1 kmol/h 
Below feed line, slope of BOL = 1.43 = L'm/V'm (liquid rate/vapor rate)

=====2. Collect or estimate the system physical properties.=====
Here it is important to know information about both the top and bottom of the column. Useful information includes temperature, pressure, column pressure drop (a common assumption is 100 mmH2O per plate), densities, molecular weights, surface tensions, and number of stages (which can be estimated from the McCabe-Thiele diagram).

=====3. Select a Trial Plate Spacing=====
The plate spacing will depend on the column diameter and operating conditions. Plate spacings from 0.15 m to 1.0 m are typically used. The smaller the diameter, the smaller the spacing. Small columns will use close spacing. Columns with diameters above 1.0 m, plate spacings of 0.3 m to 0.6 m are normally used. A good initial estimate is 0.5 m.
[[File:trayspacing.jpg|400px|center|]]

=====4. Estimate the column diameter, based on flooding considerations.=====
Vapor and liquid flow rates will vary along the column, so plate design needs to be considered both above and below the feed. Using plate spacing and FLV (which is the square root of the ratio of the liquid to vapor flow rates), you can obtain the value of K from the plot.

[[File:floodingplot.jpg|400px|center|frame|Figure. Plate Spacing]]
There is a range of vapor and liquid flow rates in which the column needs to be operated. Too low or too high of rates can result in various inefficiencies in the column operation, as shown in the figure below. For example, if the vapor rate is too high, flooding will occur. However, it is not safe to operate on the flooding line. Instead, columns are typically designed for 80% of flooding at the maximum flow rate.
[[File:vap_rate_vs_liq_rate.jpg|400px|center|frame|Figure. Tray behavior]]

=====5. Decide the liquid flow arrangement.=====
Common flow arrangements are single pass (cross flow), double pass, and reverse flow. Using conditions at the bottom of the column, calculate the max volumetric flow rate. Use this flow rate and the column diameter to determine the preferred flow arrangement from the chart below.
[[File:Liquidflow.jpg|200px|center|]]

=====6. Make a trial plate layout: downcomer area, active area, hole area, hole size, weir height.=====
Standard sizes for trays -- and good assumptions for the first iteration -- are: weir height, hw = 50mm ; hole diameter, Dh = 5mm ; plate thickness, tpl = 5mm. From the graph below, the ratio of downcomer area (Ad) to column cross-sectional area (Ac) can be determined from the ratio of weir length (lw) to column diameter (Dc) and vice versa.
[[File:platelayout.jpg|200px|center|frame|Figure. Plate Dimensions]]

=====7. Check the weeping rate=====
Compare the actual vapor velocity to the minimum vapor velocity -- if velocity is too low fluid will "weep" through the tray holes. If the weeping rate is unsatisfactory, return to step 6 and choose different values for the plate layout dimensions. From the chart in step 4, it can be seen that there is a minimum vapor flow rate below which the liquid "weeps" from the tray above.

For the remaining steps in this design process, it is recommended to check your assumptions after each step and revise them as necessary in order to maintain operation in the "sweet spot" of the vapor rate vs. liquid rate plot. Additional iterations may be required as you move through the procedure.

Calculate the maximum liquid flow rate. Calculate the minimum liquid flow rate at 70% turndown (recommended). Calculate the height over the weir as
<math>h_o=750[\frac{L_w}{p_Ll_w}]^\frac{2}{3}</math>

=====8. Check the plate pressure drop=====
<dfn>If the pressure drop calculated here is too high, return to step 6.</dfn>

Proceed to step 9 if the pressure drop assumption is valid.
=====9. Check the downcomer backup. =====
<dfn>If the downcomer backup is too high, return to step 6 or 3.</dfn>

The plate spacing affects the amount of fluid in the downcomer. Calculate the level in the downcomer and the residence time of the fluid to see if the values are valid. Note that residence times greater than 3 seconds are acceptable.

Proceed to step 10 if residence time is acceptable.
=====10. Decide plate layout details.=====
Determine calming zones, the unperforated areas at the inlet and outlet sides of the plate. The width of each zone is usually made the same. Recommended values are: below 1.5 m diameter, 75 mm; above, 100 mm. The unperforated area can be calculated from plate geometry. Also check the hole pitch, or the distance between hole centers. It should not be less than 2.0 hole diameters. A normal range is between 2.5 and 4.0 hole diameters. The shape must also be specified. Square and equilateral triangle holes are used.

=====11. Recalculate the percentage flooding based on the chosen column diameter.=====
An assumption of 80% flooding was chosen so that operation would occur in the "sweet spot." This assumption must be checked by calculating the flooding percentage for a given column diameter.
uv = (max volumetric flow rate)/(net area)

<math>%flooding = \frac{u_v}{u_f}</math>

If the hole pitch is unsatisfactory, return to step 6.
=====12. Check entrainment=====
''If too high, return to step 4'' Use the graph below to determine entrainment from FLV.
[[File:entrainment.jpg|400px|center|]]

The value for fractional entrainment can be used to re-estimate the column efficiency, and reevaluate the number of trays needed. Can return to step 1 for more accurate estimates.

=====13. Optimize design.=====
After returning to step 1 to reevaluate the number of trays, it is valuable to repeat steps 2 through 12 to find the smallest diameter and plate spacing acceptable at the lowest cost.

=====14. Finalize the design.=====
Optional: draw up the plate specification and sketch the layout of the plate.

====Bubble-Cap Trays====
Bubble-cap trays consist of a weir around each hole in the tray which is covered with a cap that has holes or slots to allow vapor passage. Entrainment is about three times larger than a sieve tray. Bubble-cap trays require larger tray spacing than sieve tray design. Bubble-cap trays have been known to have problems with coking, polymer formation, or high fouling mixtures. Recently, very few new bubble-cap columns are being built due to the expense and marginal benefits. However, engineers will likely encounter bubble-cap columns still currently in operation.

====Flow Patterns====
Cross flow columns are the most common pattern for distillation columns. For liquid flows between 50 and 500 Gal/min, a cross flow column is appropriate. When liquid flow is increased above 500 Gal/min, an engineer should consider designing a double pass or multi-pass column. This will reduce the liquid gradient on the tray and reduce the downcomer loading (Wankat, 2012).

===Column Sizing===
Column height will be dependent on the amount of trays required and the spacing between the trays. Normally, tray spacing of 0.15 m to 1 m is used. For columns, above 1 meter in diameter, 0.5 m can be used as an initial estimate.

Column diameter is influenced by the vapor flow rate in the column. The trays can not have excess liquid entrainment or high pressure drops; therefore, vapor velocity in the column must be maintained at a reasonable level.

An equation based on the Souders and Brown equation can be used as an estimate for the max allowable superficial vapor velocity,

<math>\hat u_v = (-0.171l_t^2 + 0.27l_t - 0.047){\frac{\rho_L - \rho_v}{\rho_v}}^{1/2}</math>

where <math>l_t</math> is the plate spacing in meters, <math>\rho_L</math> is the density of the liquid stream, and <math>\rho_V</math> is the density of the vapor stream.

Column diameter, <math>D_c</math>, can then be estimated using the relation,

<math>D_c = \sqrt{\frac{4\hat{V_w}}{\pi\rho_v\hat{u_v}}}</math>

where <math>\hat{V_w}</math> is the maximum vapor rate in kg/s (Towler et al., 2013).

===Distillation Applications===

Distillation is a process that can be implemented in various scales. There is both laboratory scaled distillation as well as very large industrial distillation. Other applications for distillation include food/alcohol processing and herb distillation for the perfume and medical industries. Typically laboratory scaled distillation occurs in batches whereas industrial distillation (e.g. fractional distillation of crude oil) occurs continuous with a constant distillate and bottom effluent streams.

Some applications of distillation are concerned the top stream only, some the bottom stream only and others both streams can be used for future products. In alcohol distillation for example, the water that is separated from the ethanol/water binary solution is discarded as waste water. In fractional distillation of crude oils, the heavy hydrocarbons at the bottom of the column are collected and sold along with the light hydrocarbons that appear in higher side draws (Wankat, 2012).

===Example Case: Ideal Distillation===

Assume an equimolar mixture flowing at 10 mol/s of 20 mol% n-pentane, 30 mol% n-hexane, and 50 mol% n-heptane. Separate the mixture into 3 products: 99% pure n-pentane, 99% pure n-hexane, 99% n-heptane. Assume the feed and products are all liquids at the bubble points. There are two process alternatives to consider in this example. The direct sequence removes the most volatile species, pentane, in the first column, and then separates hexane and heptane in the second column. The indirect sequence separates the heaviest product, heptane, and then separates pentane from hexane in the second column. This example will consider the direct sequence. Next, we must decide if these species exhibit fairly ideal behavior during distillation. Since the n-alkanes have very similar properties, it is safe to assume they will display close to ideal behavior. The next step is to look up the boiling points of the 3 species. In this case, the normal boiling points of pentane, hexane, and heptane are 309 K, 342 K, and 372 K, respectively. Also, it is a good idea to look up relative volatilites, to further verify near-ideality of the mixture, but also to obtain the information necessary for the Underwood method, which we will employ to obtain a solution. The next step is to write out material balances based on molar flows and the design specifications. They go as follows:

<math>\mu_I(nC5) + \mu_{II}(nC5) = 2 mol/s</math>

<math>\mu_I(nC6) + \mu_{II}(nC6) + \mu_{III}(nC6) = 3 mol/s</math>

<math>\mu_{II}(nC7) + \mu_{III}(nC7) = 5 mol/s</math>

<math>\mu_I(nC5) = 99\mu_I(nC6)</math>

<math>\mu_{II}(nC5) = (5/990)\mu_{II}(nC6)</math>

<math>\mu_{II}(nC7) = (5/990)\mu_{II}(nC6)</math>

<math>\mu_{III}(nC7) = 99\mu_{III}(nC7)</math>

where <math>\mu</math> represents the molar flow, and the subscript represents the product stream.

Solving this system of equations:

<math>\mu_I(nC5) = 1.985\ mol/s</math>

<math>\mu_{II}(nC5) = 0.015\ mol/s</math>

<math>\mu_I(nC6) = 0.020\ mol/s</math>

<math>\mu_{II}(nC6) = 2.930\ mol/s</math>

<math>\mu_{III}(nC6) = 0.050\ mol/s</math>

<math>\mu_{II}(nC7) = 0.015\ mol/s</math>

<math>\mu_{III}(nC7) = 4.985\ mol/s</math>

At this point we have enough information to use Underwood's method to estimate the minimum vapor flows in the column. The following three equations are used in Underwood's method:

<math>\sum_i \frac{\alpha_{ik}}{\alpha_{ik}-\phi}f_i = (1-q)F</math>

<math>(R_{min}+1)D = \sum_i \frac{\alpha_{ik}}{\alpha_{ik}-\phi}d_i = V_{min}</math>

<math>\bar R_{min}B = -\sum_i \frac{\alpha_{ik}}{\alpha_{ik}-\phi}b_i = \bar V_{min}</math>

where <math>\alpha_{ik}</math> is the relative volatility of species <math>i</math> to species <math>k</math>, <math>f_i</math> the molar flow of species <math>i</math> in the feed, <math>q</math> the fraction of the feed that joins the liquid stream at the feed tray, <math>F</math> the total molar flow of the feed, <math>D</math> the molar flow of the distillate, <math>R_{min}</math> the minimum reflux ratio <math>(=L_{min}/D)</math>, <math>d_i</math> the molar flow of species <math>i</math> in the distillate, <math>V_{min}</math> the minimum vapor flow possible in the top section of the column to accomplish the desired separation, <math>\bar R_{min}</math> the minimum reboil ratio <math>(=\bar V_{min}/B)</math>, <math>b_i</math> the molar flow of species <math>i</math> in the bottoms product, and <math>\bar V_{min}</math> the minimum vapor flow in the bottom section of the column. The final variable, <math>\phi</math>, will be solved for using the first Underwood equation, and it's value will be decided based on the relative volatilities of the key components in the column.

So, after solving the first Underwood equation, we get two values for <math>\phi</math>, 3.806 and 1.462. Because 3.806 is between the relative volatilities of the key components, we will substitute that value for <math>\phi</math> into the second Underwood equation. Doing so for both columns gives <math>V_{min} = 6.4\ mol/s</math> for the first column and <math>V_{min} = 8.9\ mol/s</math> for the second column, for a total minimum vapor flow of 15.3 mol/s. The process would then be repeated for the indirect sequence, and the decision for which process to use would be justified by the process with the overall minimum vapor flow (Biegler et al., 1997).

==Absorption==
===Description of Absorption===
Another separation process used in industry is absorption, which is used to remove a solute from a gas stream. It accomplishes this by contacting the gas mixture with a liquid solvent that readily absorbs the undesirable components from the gas stream, purifying the gas stream. This separation process is determined by the inputs of the liquid flow rate, temperature, and pressure.

The absorption factor, which can be determined mathematically, determines how readily a component will absorb in the liquid phase. The absorption factor of component i is

<math>A_i=L/K_iV</math>

where <math>L</math> is the liquid flow rate entering the column, <math>V</math> is the vapor flow rate entering the column, and <math>K_i</math> is the vapor/liquid equilibrium ratio for component i (Peters & Timmerhaus, 2003). Higher absorption factors result in higher absorptivity into the liquid and a decrease in the number of trays required for separation, however a diminishing return occurs after the absorption factor is greater than 2.0. An absorption factor of 1.4 is most commonly used.

In general absorption can be seperated into two overarching categories, physical and chemical absorption. In physical absorption, the unwanted solute in the gas is absorbed into the liquid phase because solubility of the component is higher in the liquid phase than the gas phase. In chemical absorption the solute is removed from the gas via a reaction with the solvent, this reacted product is then transported into the liquid phase (Danckwerts 1965). There are two types of chemical absorption reversible and irreversible. Generally reversible chemical absorption is preferred as the solvent can be put through a stripper and regenerated so it can be recycled back to the absorption process (Wankat, 2012).

===Absorption Apparatus===

There are five major apparatus used for absorption in industrial application. These five pieces of equipment are spray absorbers (or towers), ejector (venturi) scrubbers, packed columns, trayed columns, and film absorbers (Schmidt, 2012).

==== Spray Tower vs Ejector Scrubber ====

In both '''spray tower''' and the '''ejector scrubber''' nozzles are employed to produce small solvent droplets. These small droplets increase the surface area of the liquid to gas contact allowing for the maximum amount of mass transfer to occur between the gas mixture and the liquid. The major difference between the two nozzle equipment designs is the configuration and type of nozzles. In the ejector scrubber shown in Figure 3 there is a single nozzle that is generally a higher pressure spray nozzle that produces finer solvent drops allowing for an even greater amount of mass transfer enabling better physical absorption (Schmidt, 2012).
[[File:Ejectorventuri.jpg|thumb|200px|center|Figure 3. Ejector Scrubber (US EPA, 2006)]]
'''Spray towers''' on the other hand generally have many nozzle at different heights where the liquid solvent will be sprayed out of to contact the gas running through the tower. This design is used in order to ensure the gas contacts the liquid as throughout the tower. These nozzles are lower pressure than a ejector scrubbers nozzle and thus physical mixing is worse in this configuration. Since physical mixing is generally worse in this configuration it is usually used in conjunction with a chemical absorption process. The other major difference between the ejector scrubber and the spray tower is that gas and liquid flow is cocurrent in the former while it is countercurrent in a spray tower. A spray tower absorber is shown below in Figure 4 (Schmidt, 2012).
[[File:SparyTowerAbsorber.jpg|thumb|200px|center|Figure 4. Spray Tower Absorber (US EPA, 2006)]]

==== Tower Type Absorption Apparatus ====
'''Packed column absorbers''' and '''tray column absorbers''' have very high efficiencies for the removal of an unwanted solute in the gas stream. The major disadvantage a trayed column has when compared to a packed column is the pressure drop. The pressure drop in a packed column is generally very low, whereas in between each tray of a trayed column pressure drop can be quite large. However the advantages inherent to trayed columns become clear when one needs the solvent to have a high concentration of the component to be removed from the gas stream. This is most important in the case where there is a very low concentration of the component in the gas stream and the specification states the solvent must contain a high concentration of that component. In this case the flow rate of the solvent may not be high enough for a packed column, however in a trayed column the solvent flow rate can be near zero for operation (Schmidt, 2012). Packed and trayed column internals are very similar to the setups found in the respective distillation columns.

For a '''trayed column''' the plate efficiency can be calculated using O'Connell's Correlation which invovles the Henry's Law constant, total system pressure, and solvent viscosity at the operating temperature (Towler & Sinnott, 2013).

<math>x=0.062*\frac{\rho_s*P}{\mu_s*H*M_s}</math>

where
<math>x</math> is the tray efficiency,
<math>\rho_s</math> is the density of the solvent in <math>kg/m^3</math>,
<math>P</math> is the total pressure of the system in <math>N/m^2</math>,
<math>\mu_s</math> is the solvent's viscosity in <math>mNs/m^2</math>,
<math>H</math> is the Henry Law constant in <math>1/(Nm^2*(mol fraction))</math>,
and <math>M_s</math> is the molecular weight of the solvent.

A packed towers height can be determined using the equations below when concentration of solute is below 10% so that the assumption that the flow of gas and liquid will be essentially constant throughout the column holds (Towler & Sinnott, 2013). The height of packing <math>Z</math> is given by the following equation:

<math>Z=\frac{L_m}{K_G*a*P}*\int\limits_{y_2}^{y_1} \frac{dy}{y-y_e}\,</math>

where <math>P</math> is the total pressure, <math>a</math> is the interfacial surface area per unit volume, <math>y_1</math> and <math>y_2</math> are the mol fractions of the solute in the gas stream at the bottom and top of the column respectively, <math>G_m</math> is the molar gas flow rate per unit cross-sectional area, and <math>y_e</math> is the mole fraction of solute in the gas that would be in equilibrium with the liquid concentration.

The first half of the equation before the integral can be called the height of an overall gas-phase transfer unit <math>H_G</math> and the second part of the equation is the number of overall gas-phase transfer units or <math>N_G</math>. Using these definitions the above equation can be simplified to

<math>Z=H_G*N_G</math>

These equations assist in sizing an absorption column (Towler & Sinnott, 2013).

==== Film Absorber ====
The final absorber the film absorber is generally used in the case where the heat of absorption must be removed. The film absorber operates by sending the gas and solvent through a heat exchanger where the solvent creates a thin film on the walls of the tubes and the gas flows through the interior allowing for solute transfer. The good heat transfer present in a film absorber makes it preferable for situations where low temperatures are required for a high recovery of the solute (Schmidt 2012).

===Industrial Absorption Processes===
An industrial example is lean oil absorption, which is used to separate nitrogen and other impurities from natural gas. A lean oil is contacted with low quality natural gas, and the methane is selectively absorbed by the lean oil, leaving the impurities behind. The methane is subsequently regenerated from the rich oil as high quality natural gas (Petrogas Systems, 2014).

Other common industrial practices of absorption come from sour gas treatment. Amine gas treating is used to remove hydrogen sulfide or carbon dioxide from gas streams via a reversible chemical absorption. In amine gas treating the sour gas is fed to the bottom an absorber where amine solution is fed to the top along with any necessary make up water. The sour gas components are absorbed into the amine via a chemical absorption method. Sweet gas leaves the top of the absorber whereas the amine out of the bottom, now rich with acidic components is sent to a regenerator where the acid gas components are stripped and the acid gas is generally sent to a flare whereas the amine now lean again is recycled back into the first absorber (Miller & Zawacki, 1978). Figure 5 below shows the typical setup of an amine plant. Another type of sour gas treatment that uses absorption is Merichems LO-CAT process which uses a chelated iron to remove hydrogen sulfide from feed gas in the absorption column (Merichem 2015).
[[File:AmineTreating.png|thumb|400px|center|Figure 5. Amine Gas Treating Plant Schematic]]

==Stripping==
This process separates solutes from solvents (often after absorption, to purify the solvent so that it can be recycled to an absorber). Stripping will depend on the vapor and liquid flow rates, as well as the temperature and pressure of the column. There is a temperature drop down the column, so columns generally have either an increased operating temperature or decreased operating pressure.

The stripping factor of component i is

<math>S_i=K_iV/L</math>

where <math>K_i</math> is the vapor/liquid equilibrium ratio, <math>V</math> is the vapor flow rate entering the column, and <math>L</math> is the liquid flow rate entering the column, will determine how much of solute i will be stripped from the liquid into the vapor phase (Peters & Timmerhaus, 2003). The usual range for the stripping factor is between 1.2 and 2.0, with a stripping factor of 1.4 being most economic.

An example of stripping in industry is the deodorization of food items such as oils. The oil is heated and allowed to trickle down the column while steam flows up from the bottom of the column. At the vapor-liquid interface, volatile components of the oil transfer to the steam and are carried off the top of the column, leaving a purified oil product (Alfa Laval, 2014).

==Bioseparations==
===Importance===
As our ability to manipulate and engineer biological systems improves, biological products are becoming an increasingly important source of therapeutics and fuels. The production of fuels from biomass via either the enzymatic breakdown of a feedstock or the secretion of usable lipids from algae is a promising new energy source. Additionally, enzymes, antibodies and other therapeutic proteins have been applied to the treatment of a wide range of diseases. Although each process requires its own set of separations, all follow the same basic format: separation of biomass, product isolation, and product purification (Belter et al., 1998). This section will provide examples of unit operations in each step. Ultimately, the choice of separation process and unit operations will depend on the specific process and product. The descriptions below are examples of the most common bioseparation operations within the general platform (Harrison et al., 2003).

Bioprocesses begin with fermentations or growth operations. In biofuel production processes, this may involve growing algae or breaking down corn or cellulosic biomass. For the production of therapeutics, mammalian or bacterial cells may be grown in a fermentor and the product secreted into the supernatant or harvested from the cells.

===Biomass Separations===
After fermentation and product production, the solid biomass must first be separated from the desired product. If the product is secreted from the cells, this can be done immediately after fermentation ends. If the product is not secreted, the cells must first be lysed.
Cell lysis is the process of lysing, or breaking, the cell in open. Mechanical lysis is the simplest, and involves physically breaking the cell either by mashing (think mortar and pestle) or blending the cells into a homogenous solution in a homogenizer. Chemical lysis is another method, achieved by introducing an osmotic shock or chemically degrading the cell membrane. Additional separation can be achieved by flocculation, which is the process of aggregating biomaterial by charge neutralization or bridging. These larger complexes are easier to separate from smaller molecules (Harrison et al., 2003).

The next step is removing the unwanted biomass from the product in solution. Separation by centrifugation or sedimentation are the most common, although filtration is sometimes also used for processes where a biomass cake is desired. Both methods utilize density differences to separate the product from the solid biomass (Towler and Sinnott, 2013).

====Sedimentation====

Sedimentation relies purely on the force of gravity, while centrifugation speeds the settling process by subjecting the cells to a centrifugal force. Sedimentation in a settling tank is the simplest method of solid-liquid bioseparation. In this process, biomass in a tank is simply allowed to settle to the bottom over time. While this process is inexpensive, requires little energy and can separate out large volumes of biomass, it generally requires long time periods and is only mostly in very large-scale processes where active centrifugation is difficult (Belter et al., 1998).

====Centrifugation====
Centrifuges are widely utilized across many processes, and thus a wide variety of scales and designs have been developed. Disk-stack centrifuges, in which the solid phase is deposited onto “shelves” in the center of the spinner and liquid phase is pushed to the outside, are some of the most commonly used centrifuges in industry. They are especially suited to biomass separation processes because they can be built on a large scale and are ideal for separating fine solids from liquids. [[File: Disk_stack_centrifuge_towler.png|frame|center|Fig. 6: Diagram of a disk-stack centrifuge (Tolwer et al, 1997).]] Tubular bowl centrifuges are also common and can reach separation efficiencies of up to 90%. Heavier products accumulate along the sides of the bowl, while the light phase flows out the top. They separate products by can be used both to separate solids from liquids and immiscible liquids, such as and oil product and an aqueous broth (Tolwer and Sinnott, 2013). [[File: tubular bowl centrifuge towler.png|frame|center|Fig. 7: Diagram of a tubular bowl centrifuge centrifuge (Tolwer and Sinnott, 2013).]]

Centrifugation scale-up is made easier by sigma analysis, which allows for the estimation of appropriate feed rates for different size centrifuges. The sigma factor is dependent on the inner and outer radius of the centrifuge, the angular velocity, and the sedimentation velocity of the solid particles being separated. It can be thought of as the characteristic cross-sectional area with units of [length]2. The sedimentation velocity can be calculated by

<math>v_g={\frac{2a^2(\rho-\rho_0)}{9\mu}}g</math>

where <math>v_g</math> is the sedimentation velocity, <math>a</math> is the cell or biomass particle diameter, <math>\rho</math> is the particle density, <math>\rho_0</math> is the fluid density, and <math>\mu</math> is the fluid viscosity. The volumetric flow <math>Q</math> can be estimated by

<math>Q=(v_g)(\Sigma)</math>.

The equality

<math>{\frac{\Sigma_1}{\Sigma_2}}={\frac{Q_1}{Q_2}}</math>

can be an easy way to estimate equivalent flow rates between a small-scale centrifuge 1 and larger centrifuge 2 (Harrison et al., 2003).

====Example: Centrifugation Scale-up====

You are trying to separate a cell of radius 0.4 <math>\mu</math>m with a density of 1.05 g/cm3 from broth of mostly water (density of 1 g/cm3 and viscosity of 0.01 g/cm s). The sigma factor of the centrifuge you are using is 1 x 106 cm2. A] What volumetric flow rate should you use? B] If you want to scale up the process to a centrifuge with <math>\Sigma</math> = 3 x 106 cm2, what flow rate would you use in the larger centrifuge?

Solution:
A] Using the equation for <math>v_g</math>, and being mindful of units, the sedimentation velocity equals 1.74 x 106 cm/s. The flow rate, then, equals

<math>Q=(1.74 x 10^-6)(1,000,000) = 1.74 cm^3/s = 0.104 L/min</math>.

B] Keeping in mind that for the same process, <math>v_g1 = v_g2,</math> and rearranging the sigma factor equality, the new flow rate is

<math>Q_2 = {\frac{\Sigma_2 x Q_1}{\Sigma_1}} = {\frac{(3 x 10^6)(0.104)}{1 x 10^6}} = 0.313 L/min </math>

===Product Isolation===
Liquid-liquid separation, to extract the product from the aqueous phase, is much less straightforward than liquid-solid extraction. Many methods - especially adsorption, filtration, and precipitation - are similar in principle to operations found in other, non-biological separations. The exact separations used depend on the nature of the product and the scale of the process. These processes are nearly identical to their non-biological counterparts, and their description is left to other sections.

Particular care needs to be taken with protein products because of their instability, and the selection of an appropriate solvent or adsorbent is crucial to a successful process (Harrison et al., 2003).

===Product Purification===
The final steps of protein purification and polishing remove any remaining contaminants and bring the concentration of product to an appropriate value for applications. Purification processes for food-grade and medical products can be extensive, as sterility and high purity are essential. Purification in fuel-producing processes may be less extensive, depending on the process. Chromatography and crystallization are two common steps in purification and are especially used in industrial scale protein production. Several different types of chromatography exist with the ability to carry out different types of separations.
Chromatography is similar to adsorption in that it relies on differences in affinity between solutes and a solid surface. A solution is eluted through a column containing a solid resin with various affinities for the substances in solution. In adsorption, the solutes are evenly saturated throughout the column. Chromatography differs in that solutes are deposited a resin phase before the column is flushed with an elution solvent specific that results in solutes eluted in bands.

==== Ion Exchange Chromatography ====

There are two main types of ion exchange columns—anion and cation. Anion exchange resins have a positive charge and are used to retain products with a negative charge. Cation exchange resins have a negative charge and are used to retain products with a positive charge. The pH of the elution buffer is change to force a specific solute to wash out, depending on whether the pH of the buffer is above or below the isoelectric point of the solute (Belter et al., 1998). This is especially useful for the separation of protein product (including antibodies), nucleic acids, and other charged molecules. When the solutes have sufficiently different isoelectric points, the pH of the buffer is manipulated to affect the solute charge and force the product to elute while the solute remains preferentially bound to the resin, or vice versa (Harrison et al., 2003). In general, the most strongly charged molecules will remain in the column for a longer period of time. Elution washes through the weakly bound ions before the more strongly bound ions. Different speeds of elution can be visualized as in figure 8.

[[File:chromatography.png|frame|center|Fig. 8: Illustration of product bands in an elution chromatography column (Belter et al., 1998).]]

==== Size Exclusion Chromatography ====

In gel filtration chromatography, small molecules are "trapped' by the porous resin and take longer to flow through the column. Larger products will elute first because the smaller molecules are better able to penetrate the resin. This forces them to take a much longer path through the column, which means it takes longer for them to elute. This operation is often used when there is a distinct difference in size between the desired product and other solutes.

==== Affinity Separations ====

Affinity chromatography is very similar to ion exchange chromatography in that the interactions between the material in the column and the molecules in the feed. The main difference is that affinity chromatography can rely on a great variety of types of interactions. Two very common types of affinity are exploited in affinity chromatography columns. The first is immunoaffinity. Proteins are specifically bound by antibodies which can be incorporated onto beads and used in chromatography. Antibodies are designed to bind only a single protein, so these interactions are considered to be highly specific. The protein can be eluted using a buffer that changes the pH or salinity in the column, which adversely affects binding (Hage, 1999).

Proteins can be separated from very complex and unrefined mixtures using this method because of the great specificity. Antibody-protein interactions are so specific that individual proteins can be isolated from blood samples as shown in the figure below. The main drawback of this method is simply that specific interactions between proteins and binding targets do not always exist. Antibodies exist for many proteins, but for others they must be customized or created using some sort of evolution process. This is not a trivial task and can require a significant capital investment (Hage, 1999).

[[File:Affinity_Chromatography_Example.PNG|frame|center|Fig. 9: Affinity purification separating fibrinogen from human plasma using an anti-fibrinogen antibody (Hage, 1999).]]

The other main type of affinity chromatography is based on protein specific tags and the molecules or surfaces to which they bind. One of the most common types of protein tags used is the polyhistidine tag. This tag consists of 6-8 consecutive histidine residues which can be added to the exterior of the desired protein product. The addition of this tag requires alterations to the coding sequence of the protein. The polyhistidine tag binds strongly with nickel and cobalt ions. The product with the tag can then be eluted with imidazole—a small molecule with the same structure as the functional group of the amino acid histidine. Imidazole will bind the cobalt and nickel ions more strongly than the histidine in the tag. Along with chromatography, protein tag interactions can be leveraged with the use of beads that can be deposited directly into the solution containing the protein of interest (Lichty et al., 2005).

Different proteins can be purified in this manner with varying levels of efficiency. There is a very high dependence on the size of the protein, electronic properties, and steric considerations. The figure below demonstrates varying separation efficiencies with proteins of different sizes (HisPur Cobalt Resin - Thermo Fisher Scientific, 2016).

[[File:HisTagPurification.PNG|frame|center|Fig. 10: Protein gel demonstrating the separation of proteins from a sample. The far left column represents the most efficient separating conditions using cobalt and the far right column represents a negative control with no purification performed (HisPur Cobalt Resin - Thermo Fisher Scientific, 2016).]]

Several other types of tag-bead interactions can be utilized in separations processes. Maltose Binding Protein is a small protein that can be added to a protein of interest. It binds strongly with beads coated in immobilized maltose and can be released by flushing with maltose. As MBP is a full sized protein that typically must be removed from the protein of interest in order for it to be used. In this case, the site specific TEV protease is often used cleave MBP from the protein of interest. In addition, under specific circumstances, other unique tags can be used and provide varying levels of specificity in separations. The Flag tag, 3x Flag tag, Glut tag, and Strep tag. While these are all commonly used, the polyhistidine tag is the most popular because it gives the highest level of specificity. Some of the main purification tags are compared in the table below (Lichty et al., 2005).

{| class="wikitable"
|-
! Tag
! Size (aa)
! Resin
! Eluting Agent
! Capacity (mg/ml)
! Cost/10 mg
|-
| His
| 6
| Talon
| Imidazole
| 5-14
| $18
|-
| Maltose Binding Protein
| 396
| Amylose
| Maltose
| 3
| $12
|-
| Glutathione S-Transferase
| 218
| GSH-Sepharose
| Glutathione
| 10
| $36
|-
| Strep II
| 8
| Strep-Tactin Sepharose
| Desthiobiotin
| 0.1
| $293
|-
| FLAG
| 8
| Anti-FLAG M2 MAb Agarose
| Flag peptide
| 0.6
| $1045
|}

Unfortunately, based on the contents of the solution, the characteristics of the protein, and the specific product being used, there are extremely large amounts of variation in yield and purity for each type of tag. Therefore, it is crucial to allocate time towards preliminary tests and scale-up experiments with the exact solution from which the protein of interest is being separated. The his tag is typically the most used because, as seen in the table above, it has one of the lowest costs and highest yields. In addition, because it is only a 6 amino acids, it often is not cleaved from the protein of interest which minimizes the number of steps required in the overall purification process (Lichty et al., 2005).

==== Crystallization ====

Crystallization, or the formation of solute crystals from a solution, is especially useful in biomolecule separations because it is possible to obtain a 99.9%+ product purity. In crystallization, a diluent is added to the homogeneous solution that reduces the solubility of the product to the point that it “falls out” of solution and crystallizes. It is similar to precipitation but results in the formation of crystals rather than unordered aggregates. Crystallization can be used on a laboratory scale for determining protein structure, on on the industrial scale for antibody and therapeutic protein productions. Batch crystallizers are often used in industry because of their simplicity and inexpensiveness compared to continuous crystallization (Harrison et al., 2003).

==Membrane Separation==
Membrane separation takes advantage of the selective permeability of membranes; they allow certain particles to pass through and selectively stop other, generally unwanted, particles. The component that passes through is called the permeate and the component stream that is rejected is called the retentate or concentrate. The applicability of membranes comes from the fact that their selectivity is determined by their pore size, which can be controlled during the creation of the membranes. Additionally, Membrane processes do not require heat meaning they generally require less energy than conventional separations technology such as distillation and crystallization. Membrane separations processes are generally classified as microfiltration, ultrafiltration, or nanofiltration depending on the size of the particles to be filtered out.

[[File:membrane techs.png|frame|Figure. Cutoffs for different membrane categories]]

===Membrane Selection, Construction, and Flow Geometries===
Membrane permeability and selectivity are the two most important factors to consider when selecting a membrane. For gas separations, the permeation of the gas is usually facilitated by the gas dissolving in the membrane on one side and then evaporating on the permeate side. Therefore permeability depend largely on the solubility of components in the membrane.

The two most commonly utilized membrane configurations are hollow fiber and spiral wound. Hollow fiber is generally the most commonly utilized module for gas separations. These are formed by gluing the two ends of the hollow fiber to a resin forming a closure. The fibers are housed in a shell much like a heat exchanger. The feed flows past thousands of tubes with the permeate flowing into the hollow tubes and out the closure. The retentate then flows out of the shell not having gotten through the membrane.

[[File:hollow membrane.jpg|frame|Figure. Hollow fiber membrane module]]

Spiral wound membranes are created by sealing two membrane sheets back to back on three edges to form a sort of pocket. This fourth open edge is then attached to a porous tube which allows permeate to go through it. Several membrane pockets are attached to a single tube and wrapped around in a spiral.

[[File:spiral membrane.jpg|frame|Figure. Spiral wound membrane module]]

Flow geometry is usually either dead-end geometry or cross flow geometry. In dead end, the fluid flow is normal to the membrane surface while cross flow is parallel to the membrane surface. Dead end geometry is usually used with hollow fiber membranes while cross flow is used with spiral wound membranes. Each geometry has advantages and disadvantages. Dead end geometry is generally cheaper to set up and therefore has lower initial capital costs. However, it is very vulnerable to membrane fouling, which reduces the effectiveness of the membrane. This is usually the geometry set up for small scale lab experiments. The tangential flow devices are more cost and labor-intensive, but they are less susceptible to fouling due to the sweeping effects and high shear rates of the passing flow. Most commercial industrial membrane separations are done using spiral wound cross flow membrane modules.

===Applications===
====Food Industry====
Due to the fact that MD can be conducted at relatively low feed temperatures, it was successfully tested in many areas where high temperature applications lead to degradation of the process fluids especially in food processing. It was demonstrated that MD can be used for the concentration of milk, for the recovery of volatile aroma compounds from black currant juice, and for the concentration of many other types of juices including orange juice, mandarin juice, apple juice, sugarcane juice, etc.
====Reverse Osmosis====
Reverse osmosis is the most widely used membrane separation process. In this process, fresh water passes through the membrane while dissolved salts and other solids are rejected and stay in the concentrate. In this process, feed water is pressurized in order to overcome the osmotic potential difference between the salty retentate and the fresh water desired. These processes are generally run using spiral wound membrane cylinders using a cross flow setup.

===Membrane Model===
The two most important components when considering different membranes are the permeability, which will determine flux through the membrane, and selectivity, which will determine what passes through the membrane and how much. The flux through a membrane is defined as:
<math> M_i = \frac{P_i}{δ}(p_{i,f} - p_{i,p})</math>
Where Mi is the molar flux of component i, Pi is the permeability of the membrane for component i, δ is the membrane thickness, and pi,f and pi,p are the partial pressures of component i on the feed side and permeate side respectively.
The average flux across a long cylindrical membrane such as the spiral wound module is given by:
<math> \int_0^Lm \frac{M_i,dx}{L_m}</math>
Where Lm is the length of the cylinder and x is length in meters

Membrane selectivity of the ideal separation factor is given as the ratio of the permeability of one substance over another as shown:

<math> S_(i,j) = P_i/P_j </math>

where Sij is the selectivity of the membrane for component i over j.

==Cyclones==
Centrifugal Separators, more commonly know as Cyclones, are one of the most widely used gas-solid separators. They are typically used for particles that are 5μm or larger, but if agglomeration occurs they can sometimes separate particles as small as particles as small as .5μm. These cyclones can be designed for high efficiency separations or for high throughput (Towler, 2012).

[[File:cyclone.png|thumb|center|Reverse-flow Cyclone (Towler, 2008)|500x300px]]

===Design===
Design Procedure
There are multiple design methods for cyclones. The deign method discussed below is the Stairmand method. This involved creating two designs one of which was for high throughput and the other which was designed for high efficiency. Using these base designs future cyclones can simply be estimated using scaling factors (Towler, 2012).

1. Select whether efficiency or performance is more important for your design

2. Determine the particle size distribution in your stream

3. Estimate the number of cyclones needed

3. Calculate the cyclone diameter

4. Calculate the scale up factor

5. Calculate the cyclone performance and overall efficiency

6. Calculate the cyclone pressure drop

7. Cost the system

====Particle Size====
The diameter of particles separated by cyclones are governed by this design equation. This equation is dependent upon the standards of the cyclone and how much you are trying to deviate from its standard operating conditions. Looking into the equation it shows that a larger internal cyclone diameter, higher flowrate, higher difference in density, and lower viscosity result in smaller particles being seperated. The whole term after d1 is known as a scale up factor. When designing a cyclone you would determine the flow rate, viscosity, particle size, and density difference would be known prior to using this design equation, so this would be used to optimize the diameter of the cyclone you were designing (Towler, 2012).

[[File:Towler 10.8.PNG|center|Reverse-flow Cyclone (Towler, 2012)|500x300px]]

d1=mean diameter of particle separated at the standard conditions, at the chosen
separating efficiency

d2= mean diameter of the particle separated in the proposed design, at the same
separating efficiency

Dc1 =diameter of the standard cyclone

Dc2 =diameter of proposed cyclone

Q1 =standard flow rate

Q2 =proposed flow rate

Δρ1 =solid-fluid density difference in standard conditions

Δρ2 =density difference, proposed design

μ 1=test fluid viscosity

μ 2=viscosity, proposed fluid.

[[File:Towler10.44.PNG||thumb|center|(a) Standard Cyclone Dimensions (b) High Gas Rate Cyclone (Towler, 2012)|500x300px]]

====Efficiency====
Efficiency is the percentage of total solids that are removed from the stream The graphs below were experimentally for the two standard designs Stairmand Method. This graph has the efficiency vs the particle size (Towler, 2012).

[[File:Towler10.45.PNG||thumb|center|Performance curves, standard conditions. (a) High-efficiency cyclone performance curves, standard conditions. (b) High gas-rate cyclone (Towler, 2008)|600x400px]]

If you have to scale the efficiency it can be estimated by the point where a horizontal line from the efficiency of the original particle size intersects the new particle size. A demonstration is on the graph below.

[[File:Towler 10.46.PNG||thumb|center|Scaled Performance Curve(Towler, 2008)|500x300px]]

====Pressure Drop====
One of the most important design factors in a cyclone is its pressure drop. The pressure drop occurs because entry and exit losses, kinetic energy losses, and friction losses that occur inside the cyclone. The pressure drop can be calculated using the following equation (Towler, 2012).

[[File:Towler 10.9.PNG|center|Reverse-flow Cyclone (Towler, 2008)|500x300px]]

The variables are defined as follows:
DP = cyclone pressure drop, millibars;
rf = gas density, kg/m3
;
u1 = inlet duct velocity

u2 = exit duct velocity

rt = radius of circle to which the center line of the inlet is tangential

re = radius of exit pipe

φ =factor from figure below

ψ = fc(As/A1)

fc = friction factor

As = surface area of cyclone exposed to the spinning fluid

A1 = area of inlet duct

[[File:Towler_10.47.PNG|thumb|center|Cyclone Pressure Drop Factor (Towler, 2008)|600x400px]]

====Cost Estimation====
The following values are represented in 2002 dollars. Cost is mostly a function of the amount of gas that need to be passed through the cyclone and the units are in standard meters cubed per second (sm3/sec) (Towler, 2012).

The costs can be broken down into three categories:

Capital Costs: $4600-$7600 sm3/sec

Operations and Maintenance: $1500-$18,000 sm3/sec

Annualized Cost: $2800-$29000 sm3/sec

This all comes to $.47-$440 per metric ton of pollutant removed. The high variation in these costs have to do with smaller units being much less efficient per unit of pollutant removed. Furthermore the cost is also effected by the amount of particulates that are in the gas to begin with. For example a gas with 10 g/sm3 and air with 100sm g/sm3 would have similar pumping costs, but very different cost per unit of solids removed (EPA).

==Other Separation Processes==
===Extraction===
Liquid-liquid extraction is a process for components with overlapping boiling points and azeotropes. The process requires a solvent such that some of the components of the mixture are soluble, and then the components will be separated based on this solubility in the liquid. This process can operate at moderate temperatures and pressures, so is not very energy intensive. However, a distillation column is required to extract the solvent for recycle. More recently, supercritical fluids have replaced liquid solvents in some processes for L/L extraction, due to the solute’s ability to more rapidly diffuse through them. The issue with these fluids, however, is that they must be operated at extremely high pressures and temperatures, increasing both capital and operating expenses of the process (Peters & Timmerhaus, 2003).

===Crystallization===
This process recovers solutes that have been dissolved in solution. The resulting product is in the solid phase. Depending on the material properties of the solute and solvent, the solute is recovered by precipitation after cooling, removal of solvent, or adding precipitating agents. Crystallizers are designed based on phase equilibria, solubilities, rates and amounts of nuclei generated, and rates of crystal growth. Every crystallization process is a unique system, so plant evaluation is usually required before complete implementation. Crystallization can be performed in both batch and continuous processes, and design features can control crystal size to an extent (Peters & Timmerhaus, 2003).

===Membrane Separation===
This separations process uses selectively permeable membranes to separate components in a mixture. Typically, one of the components will freely pass through the barrier while the other components will not. The stream that passes through the membrane is the permeate and the stream that does not pass is the retentate. The driving force behind this separation is a pressure gradient. Membrane separation is beneficial because it can separate mixtures at the molecular and small particle level. Furthermore, there is no phase change required so the energy input is low. Limitations of this process include achieving high product purity, incompatibility with certain stream components, low operating temperature, and low flow rates. Although membrane separation is generally not scaled up, examples of scaled-up membrane separation include seawater desalination and hydrogen recovery (Peters & Timmerhaus, 2003).

===Adsorption===
Adsorption involves an adsorbent and adsorbate. The adsorbent is typically a solid, and will typically separate the adsorbate from the stream. This process usually includes a desorption step that regenerates the adsorbent for further use. Raising the temperature or increasing the concentration of the adsorbate can reverse the adsorption process. Although the recycle of the adsorbent is a very economic design feature, the downside of this step is that it results in a cyclic process, which introduces complexity to the overall process. Industrial applications of this process are for bulk separations and gas purification. The adsorption/desorption process in these situations involves a large amount of heat transfer, which design engineers must take into account when sizing and selecting equipment material (Peters & Timmerhaus, 2003).

===External Field/Gradient Separation===
These separations use external force fields or temperature gradients to separate responsive molecules or ions. The use of these processes is fairly limited to a few specialized industrial applications (Peters & Timmerhaus, 2003).

===Settling and Sedimentation===
In settling processes, solid particles or liquid drops are separated from a stream by gravity. The stream can be in either the liquid or gas phase. For vapor-liquid mixtures, flash drums are generally used to separate the mixture. The velocity of the vapor must be less than the settling velocity of the liquid drops for this separation to occur. For liquid-liquid separation, the horizontal velocity of the fluid must be low enough to allow the low-density droplets to rise to the interface and the high-density droplets to move away from the interface and coalesce. In sedimentation, the result of the process is a more concentrated slurry. Typically a flocculating agent is used to aid in the settling process. One way to perform this separation is to use a cone-shaped tank with a slowly revolving rake that scrapes and moves the thickened slurry to the center of the cone for removal (Peters & Timmerhaus, 2003).

[[File:Example.jpg]]

====Clarifiers====
[[File:Circular_Clarifier.png|300px|thumb|bottom|Figure 9: Circular clarifier with some components labelled.]] [[File:Rectangular_Clarifier.png|300px|thumb|bottom|Figure 10: Rectangular clarifier with some components labelled.]]

Clarifiers are one of the methods used for the continuous removal of particulate solids from liquids through sedimentation by gravity. Applications include process water pretreatment, waste water treatment, and drinking water purification. Historically, clarifiers were originally developed to limit nutrient input into surface water due to fear of eutrophication. Today, they have a number of uses, particularly in wastewater treatment processes, metal removal, disinfection, and membrane pretreatment. The process helps removed dissolved solids, silt, and undesirable metals from the water, making it more suitable for downstream processes as well as human consumption (Wilson, 2005).

Clarifiers are typically used in conjunction with coagulation or flocculation agents, which promote dissolved particles to join into clumps and settle out of solution (Towler and Sinnot, 2012). Clarifiers typically consist of a large circular tank with a rotating rake at the base which scrapes settled solids towards the center. In the case of a rectangular clarifier, they are scraped to one side. Diagrams of both are represented in figures 9 and 10, respectively (NMED Surface Water Quality Bureau, 2015). Separated solids are allowed to settle to the bottom of the tank as a sludge, whereupon they are collected by the rake and disposed of properly. In the case of floating contaminants, it is possible for the clarifier to include a skimmer as well.

Clarifier efficiency varies with certain factors, including the settling characteristics of solids removed and the surface overflow rate of the tank. Clarifier efficiency can be found using the following relation:

<math>\begin{align}
E_{TSS} &= E_{TSSmax}\left ( 1 - e^\frac{\lambda}{SOR} \right )
\end{align}</math>

where <math>E_{TSS}</math> is the efficiency of total suspended solids (TSS) removal, <math>E_{TSSmax}</math> is the maximum possible efficiency, <math>\lambda \left [\frac{m}{d} \right ]</math> is the settling constant, and <math>SOR \left [\frac{m^3}{m^2 d} \right ]</math> is the surface overflow rate. The effect of flocculation chemicals on TSS can be seen in figure 11. However, it should be noted that chemical addition will increase sludge quantity and may have an adverse effect on plant aesthetics, which increases maintenance costs (Wilson, 2005).

[[File:Chem_Addition.png|200px|thumb|bottom|Figure 11: The effect of flocculating agents on total suspended solids removal in clarifiers.]]

=====Lamella Clarifiers=====

Lamella clarifiers use inclined plates in order to maximize the settling area for solids. Solids continue to settle into a hopper at the bottom of the tank while clarified water exits up through the inclined plates. This allows for the design of a smaller tank, which leads to large savings in capital costs. A lamella clarifier is pictured in figure 12.

[[File:Lamella_Clarifier.png|300px|thumb|bottom|Figure 12: A lamella clarifier with components labeled.]]

Typically, inclined plates are installed at an angle of 45 to 60 degrees and spaced 40 to 120 mm apart, which increases effective settling surface area by a factor of 6 to 12 compared to traditional clarifiers. For effective use, it is recommended that the Reynolds number be below 2000, Froude number higher than 10-5,and detention time be longer than 3 to 5 minutes. For this implementation, the equations are as follows:

<math>\begin{align}
N_{Re} &= \frac{VR}{\nu} \\
N_{Fr} &= \frac{V^2}{Rg}
\end{align}</math>

where <math>R</math> refers to the hydraulic radius, which is the cross-sectional area of the lamella, <math>V</math> is the liquid velocity, <math>\nu</math> is the kinematic viscosity, and <math>g</math> is the gravitational constant (Wilson, 2005).

=====Advantages=====

Clarifiers offer a proven, relatively inexpensive solution for solids removal. The chemical coagulants used are cheap and provide a low operating cost as well as simple maintenance. Construction is typically simple, leading to low capital costs and equipment that is easy to accommodate and maintain. Their design is also flexible, with various options such as skimmers and scrapers offering increased removal efficiency (Wilson, 2005). Operation of clarifier tanks also has lower energy requirements than membrane filtration for solids removal, given that most of the separation is aided by gravity. Water exiting clarifier units has a silt density index (SDI) averaging 4.0, which is low enough for further membrane treatment such as reverse osmosis (Prihasto, 2009).

=====Disadvantages=====

Clarifiers necessitate low turbulence to prevent resuspension of solids. This essentially requires a low entrance velocity, which can limit the production rate of certain processes or call for more clarifier units, which would drive up costs. Furthermore, clarifiers require frequent cleaning before sludge becomes too difficult to remove and reduces effectiveness. In the case of lamella clarifiers, sludge buildup on the inclined plates results in uneven flow distribution which could harm efficiency (US EPA, 2003). For this reason, maintenance requirements for lamella clarifiers are higher, but they can be reduced through the implementation of removable plates (Wilson, 2005). Clarifiers also only remove solids, so pH will not be affected, leading to the need for further pH adjustment (NMED Surface Water Quality Bureau, 2015).

=====Clarifier Design Calculations and Typical Design Values=====

======Detention Time======

Detention time (DT) is the time is takes for a unit of water to travel from the inlet of the clarifier unit to the outlet. During typical operations, the design value for this is 2 to 3 hours.

<math>\begin{align}
DT &= \frac{Tank\ Volume}{Influent\ Rate}
\end{align}</math>

======Surface Overflow Rate======

Surface overflow rate (SOR) measures the flow into the clarifier per square foot of surface area. Typical design values are 400 to 800 gal/day/sq. ft.

<math>\begin{align}
SOR &= \frac{Volumetric\ Flow\ Rate}{Surface\ Area}
\end{align}</math>

======Weir Overflow Rate======

Weir overflow rate (WOR) describes the flow in gallons per day per linear foot of weir. Typical values are 10,000 gal/day/ft.

<math>\begin{align}
WOR &= \frac{Volumetric\ Flow\ Rate}{Weir\ Length}
\end{align}</math>

======Solids Loading Rate======

Solids loading rate (SLR) describes the mass of solids in the clarifier influent per square foot of surface area. This value should not exceed 30 lbs/day/sq. ft.

<math>\begin{align}
SLR &= \frac{Solids\ Mass\ Flow\ Rate}{Surface\ Area}
\end{align}</math>

===Flotation===
Flotation is a process designed for specific solid-solid mixtures. It works by generating gas bubbles in a liquid that attach to selected solid particle. Afterwards, the particles rise to the liquid surface where they are removed by an overflow weir or mechanical scraper. The separation depends on the surface properties of the particles and its preference to attach to the gas bubbles. To meet the necessary requirements of the flotation process, a number of additives can be used to control things like the pH of the liquid-solid mixture, the activity of the solid surface, and the froth that can assist in separation. The bubbles can be produced by gaseous dispersion, dissolution, or electrolysis of the liquid (Peters & Timmerhaus, 2003).

===Centrifugation===
This process is similar to external field separation in that an external force field is applied to separate a mixture. When gravity separation is too slow due to particle densities, particle size, settling velocity, or the formation of an emulsion, centrifugation is commonly used. Centrifugal force increases the total force acting on the particle and results in faster separation times. This process is generally used to separate solids from liquids, however it can also be used to separate two liquids with very different densities (Peters & Timmerhaus, 2003).

===Drying===
Drying is performed to remove liquid from a liquid-solid mixture and produce a dry solid. Water is most often the liquid removed, but organic liquids are removed from solids on occasion as well. The heat required to vaporize the liquid is usually obtained by a series of gas-solid contacting devices. Feed condition and temperature sensitivity of the solid dictate the type of contacting device that is used. There are two groups of dryers that differ by the dependence of either mechanical means or fluid motion for gas solid contact. Another feature of dryers is to use either direct (hot gas) or indirect (conductive surface) heating (Peters & Timmerhaus, 2003).

===Evaporation===
Evaporators separate solvents from a solution by evaporation. The difference between evaporation and distillation is that evaporation requires the solute be nonvolatile. Because of this, a high separation can be achieved with one stage. Evaporators are essentially reboilers, so evaporation is a very energy-intensive process with a high thermal economy (Peters & Timmerhaus, 2003).

===Filtration===
Filtration is a process that separates a mixture of solid in a liquid or gas by passing the mixture through a porous medium in which the particles do not pass. Filtration is done by either cake filtration (particles found on the surface of the filter) or depth filtration (particles found within the filter). Cake filtration is generally performed with a cloth as the filtration medium (Peters & Timmerhaus, 2003).

==Conclusion==
Separation is a key part of most chemical processes, and there is a great variety of techniques to perform separation of compounds based on size, volatility, charge, and many other features. A common technique with which the process engineer should be familiar is distillation, but he or she should also be aware of the other available options. Some techniques may be less expensive, less energy-intensive, or more effective than distillation, depending on the specific separation problem. Therefore, the separation strategy should be carefully considered.

==References==
Air Control Technology Fact Sheet. Enviornmental Technology Fact Sheet. http://www3.epa.gov/ttncatc1/dir1/fcyclon.pdf

Belter PA, Cussler EL, Hu WS. Bioseparations: Downstream Processing for BIotechnology. New York: John Wiley; 1998.

Biegler LT, Grossmann IE, Westerberg AW. Systematic Methods of Chemical Process Design. Upper Saddle River: Prentice Hall; 1997.

Danckwerts P (1965) The Absorption of Gases in Liquids. Pure and Applied Chemistry UK 10:625-642.

Development Document for the Final Effluent Limitations Guidelines and Standards for the Metal Products and Machinery Point Source Category (Report). US Environmental Protection Agency. 2003.

Erwin, D. Industrial Chemical Process Design. New York: McGraw Hill, Professional Engineering; 2002.

Hage, David S. 1999Affinity Chromatography: A Review of Clinical Applications. Clinical Chemistry 45(5): 593–615.

Harrison RG, Todd P, Rudge SR, Petrides, DP. Bioseparations Science and Engineering. New York: Oxford University Press; 2003.

Hage, David S. 1999Affinity Chromatography: A Review of Clinical Applications. Clinical Chemistry 45(5): 593–615.

HisPur Cobalt Resin - Thermo Fisher Scientific N.d. https://www.thermofisher.com/order/catalog/product/89964, accessed February 21, 2016.

Lamella Plate Clarifier. Hydro International Web site. Available at: http://www.hydro-int.com/uk/products/lamella-plate-clarifier?s=0&r=uk. Accessed February 2, 2016.

Lean Oil Absorption. PetroGas Systems Web site. Available at: http://petrogassystems.com/technology/natural-gas-processing-and-dew-point-control/lean-oil-absorption. Accessed February 19, 2014.

Lichty, Jordan J., Joshua L. Malecki, Heather D. Agnew, Daniel J. Michelson-Horowitz, and Song Tan 2005Comparison of Affinity Tags for Protein Purification. Protein Expression and Purification 41(1): 98–105.

Merichem Gas Technologies. ®LO-CAT PROCESS available at http://www.merichem.com/images/casestudies/Desulfurization.pdf Accessed 6 Feb. 2015.

Miller L.N. & Zawacki T.S. , US 4080424, "Process for acid gas removal from gaseous mixtures", issued 21 Mar 1978, assigned to Institute of Gas Technology

NMED Surface Water Quality Bureau, New Mexico Water Systems Operator Certification Study Manual, New Mexico Environment Department, 2015.

Peters MS, Timmerhaus KD. Plant Design and Economics for Chemical Engineers. 5th ed. New York: McGraw Hill; 2003.

Prihasto, N; Lui, Q; Kim, S. Pre-treatment strategies for seawater desalination by reverse osmosis system. 2009; 249(1): 308-316. doi:10.1016/j.desal.2008.09.010

Schmidt Eberhard (2012) Waste Gases, Separation and Purification. Ullman’s Encyclopedia of Industrial Chemistry Germany 2:174-181.

Seider, W.D., Seader, J.D., and Lewin, D.R. (2004). ''Process Design Principles: Synthesis, Analysis, and Evaluation.'' New York: Wiley.

Stripping Column. Alfa Laval Web site. Available at: http://www.alfalaval.com/solution-finder/products/soft-column/Documents/Stripping%20Column.pdf. Accessed February 19, 2014.

Towler G, Sinnott R. Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013.

Turton, R.T., Bailie, R.C., Whiting, W.B., and Shaewitz, J.A. (2003). ''Analysis, Synthesis, and Design of Chemical Processes'' Upper Saddle River: Prentice-Hall.

Wankat, P.C. (2012). ''Separation Process Engineering.'' Upper Saddle River: Prentice-Hall.

Wilson, T.E., Clarifier Design, 2nd Ed., McGraw-Hill: New York, 2005.

Separation processes

2016-02-22T01:30:59Z

Bts511: /* Cost Estimation */

Authors: Nick Pinkerton, [2014] Karen Schmidt, [2014] James Xamplas, [2014] Emm Fulk, [2015] and Erik Zuehlke, [2015] John Dombrowski [2016] , Brett Sleyster [2016] , Robert Cignoni [2016] , and Osman Jamil [2016] 

Stewards: David Chen, Jian Gong, and Fengqi You

Date Presented: February 9, 2014 /Date Revised: February 1, 2014

 

==Introduction==
Essentially all chemical processes require the presence of a separation stage. Most chemical plants comprise of a reactor surrounded by many separators. Separators have a countless number of jobs inside of a chemical plant. A separator can process raw materials prior to the reaction, remove incondensable gases, remove undesired side products, purify a product stream, recycle materials back into the process, and many other jobs that are essential to the process.

Chemical engineers must understand the science of separation and the variety of ways that separation can take place. There are many ways to perform a separation some of these including: distillation, absorption, stripping, and extraction. The science of separation revolves around the presence of two phases that are in contact and equilibrium (Wankat, 2012).

[[File:Sepmeth.JPG|frame|Figure 1. Separation methods by property]]

==Theory==
===Vapor-Liquid Equilibrium===
Separation processes are based on the theory of vapor-liquid equilibrium. This theory states that streams leaving a stage in a separation process are in equilibrium with one another. The idea of equilibrium revolves around the idea that when there is vapor and liquid in contact with one another they are in constantly vaporizing and condensing. Different components in the mixture will condense and vaporize at different rates. There are three types of equilibrium conditions that can be subdivided into thermal, mechanical and chemical potential categories. These separate equilibrium states are given as:

<math>T_{liquid} = T_{vapor}</math>

<math>p_{liquid} = p_{vapor}</math>

<math>chemical potential_{liquid} = chemical potential_{vapor}</math>

==Distillation==
===Flash Distillation===
Flash Distillation is one of the simpler separation processes to be employed in a chemical plant. The main premise of flash distillation is that a portion of a liquid feed stream vaporizes in a flash chamber or a vapor feed condenses. Vapor-liquid equilibrium will cause the vapor phase and the liquid phase to have different compositions. The more volatile component of the mixture will compose of a larger portion of the vapor. This simple separation is easy to manufacture but does not result in large degrees of separation.

Flash distillation requires a feed stream that is pressurized and heated and then passed through a valve into a flash drum. The large pressure drop across the valve will result in a partial vaporization of the fluid. Vapor will be removed overhead from the flash drum while the remaining liquid will collect at the bottom of the drum and be removed. Most flash drums will contain an entrainment eliminator which is a screen that prevents liquid from being carried into the vapor effluent. Figure 2 shows a simple overview of the flash distillation process. As shown, there is a heater that flows into a let-down valve where the two-phase flow begins. Variables y and x are the mole fractions of the more volatile component in the vapor and liquid effluents, respectively.

[[File:Flash.gif|center|frame|Figure 2. Flash Distillation Flow Diagram]]

===Column Distillation===
Distillation columns are the most widely used separation technique used in the chemical industry, accounting for approximately 90% of all separations (Wankat, 2012). Distillations in columns consist of multiple trays that each act at their own equilibrium conditions. Large columns are able to perform complete separations of binary mixtures as well as more complex multi-component mixtures.

[[File:column.jpg|250px|center|]]
===Stages===
Columns are separated into stages by the presence of trays. These trays allow for vapor-liquid contact and equilibrium to occur. Typically, the more stages in a column, the larger separation that can be achieved. There are many different types of trays that can be used in a column.
====Sieve Trays====
The simplest and least expensive tray type is the sieve tray which is a sheet of metal with holes punched into it to allow vapor flow. Sieve trays can have different hole patterns and sizes that will affect the tray efficiency and flow rates.

[[File:sieve.jpg|200px|center|]]

====Sieve Tray Design Procedure====

The design of these plates is done through a trial-and-error process. Most commercial process simulations (such as HYSYS) have default tray designs, and automatically specify dimensions. However, these dimensions selected or calculated by the simulations may not give the best performance for your system, so it is valuable to understand how to design the sieve trays and how specific parameters may affect performance. Hand calculations using the following methods can be used to guide the simulation programs to better design. This section will use sample data to work through an example of the process. The following is a general list of steps for designing a sieve plate:

=====1. Calculate the maximum and minimum vapor and liquid flow rates for the turndown ratio required.=====
This data can be collected from a McCabe-Thiele diagram and/or from process simulation data.

Data from McCabe Thiele diagram, for example: 
Number of stages = 10 
Slope of top operating line = 0.185 
Slope of bottom operating line = 1.43 
Top composition = 98.8 mol% acetone 
Bottom composition = 4 mol% acetone (a) 
Minimum reflux ratio = 0.31 

Collect data rom mass balances: 
MW of feed = (0.6 kmol a/kmol)*(58 kg/kmol a) + (0.4 kmol acetic acid/kmol)*(60 kg/kmol aa) = 59 kg/kmol 
Feed rate = 100 kmol/h 
Using material balances... D = 57.3 kmol/h ; B = 42.7 kmol/h 
Vapor rate, V = D(1+R) = 57.3 kmol/h (1+.31) = 75.1 kmol/h

=====2. Collect or estimate the system physical properties.=====
Here it is important to know information about both the top and bottom of the column. Useful information includes temperature, pressure, column pressure drop (a common assumption is 100 mmH2O per plate), densities, molecular weights, surface tensions, and number of stages (which can be estimated from the McCabe-Thiele diagram).

=====3. Select a Trial Plate Spacing=====
The plate spacing will depend on the column diameter and operating conditions. Plate spacings from 0.15 m to 1.0 m are typically used. The smaller the diameter, the smaller the spacing. Small columns will use close spacing. Columns with diameters above 1.0 m, plate spacings of 0.3 m to 0.6 m are normally used. A good initial estimate is 0.5 m.
[[File:trayspacing.jpg|400px|center|]]

=====4. Estimate the column diameter, based on flooding considerations.=====
Vapor and liquid flow rates will vary along the column, so plate design needs to be considered both above and below the feed. Using plate spacing and FLV (which is the square root of the ratio of the liquid to vapor flow rates), you can obtain the value of K from the plot.

[[File:floodingplot.jpg|400px|center|frame|Figure. Plate Spacing]]
There is a range of vapor and liquid flow rates in which the column needs to be operated. Too low or too high of rates can result in various inefficiencies in the column operation, as shown in the figure below. For example, if the vapor rate is too high, flooding will occur. However, it is not safe to operate on the flooding line. Instead, columns are typically designed for 80% of flooding at the maximum flow rate.
[[File:vap_rate_vs_liq_rate.jpg|400px|center|frame|Figure. Tray behavior]]

=====5. Decide the liquid flow arrangement.=====
Common flow arrangements are single pass (cross flow), double pass, and reverse flow. Using conditions at the bottom of the column, calculate the max volumetric flow rate. Use this flow rate and the column diameter to determine the preferred flow arrangement from the chart below.
[[File:Liquidflow.jpg|200px|center|]]

=====6. Make a trial plate layout: downcomer area, active area, hole area, hole size, weir height.=====
Standard sizes for trays -- and good assumptions for the first iteration -- are: weir height, hw = 50mm ; hole diameter, Dh = 5mm ; plate thickness, tpl = 5mm. From the graph below, the ratio of downcomer area (Ad) to column cross-sectional area (Ac) can be determined from the ratio of weir length (lw) to column diameter (Dc) and vice versa.
[[File:platelayout.jpg|200px|center|frame|Figure. Plate Dimensions]]

=====7. Check the weeping rate=====
Compare the actual vapor velocity to the minimum vapor velocity -- if velocity is too low fluid will "weep" through the tray holes. If the weeping rate is unsatisfactory, return to step 6 and choose different values for the plate layout dimensions. From the chart in step 4, it can be seen that there is a minimum vapor flow rate below which the liquid "weeps" from the tray above.

For the remaining steps in this design process, it is recommended to check your assumptions after each step and revise them as necessary in order to maintain operation in the "sweet spot" of the vapor rate vs. liquid rate plot. Additional iterations may be required as you move through the procedure.

Calculate the maximum liquid flow rate. Calculate the minimum liquid flow rate at 70% turndown (recommended). Calculate the height over the weir as
<math>h_o=750[\frac{L_w}{p_Ll_w}]^\frac{2}{3}</math>

=====8. Check the plate pressure drop=====
<dfn>If the pressure drop calculated here is too high, return to step 6.</dfn>

Proceed to step 9 if the pressure drop assumption is valid.
=====9. Check the downcomer backup. =====
<dfn>If the downcomer backup is too high, return to step 6 or 3.</dfn>

The plate spacing affects the amount of fluid in the downcomer. Calculate the level in the downcomer and the residence time of the fluid to see if the values are valid. Note that residence times greater than 3 seconds are acceptable.

Proceed to step 10 if residence time is acceptable.
=====10. Decide plate layout details.=====
Determine calming zones, the unperforated areas at the inlet and outlet sides of the plate. The width of each zone is usually made the same. Recommended values are: below 1.5 m diameter, 75 mm; above, 100 mm. The unperforated area can be calculated from plate geometry. Also check the hole pitch, or the distance between hole centers. It should not be less than 2.0 hole diameters. A normal range is between 2.5 and 4.0 hole diameters. The shape must also be specified. Square and equilateral triangle holes are used.

=====11. Recalculate the percentage flooding based on the chosen column diameter.=====
An assumption of 80% flooding was chosen so that operation would occur in the "sweet spot." This assumption must be checked by calculating the flooding percentage for a given column diameter.
uv = (max volumetric flow rate)/(net area)

<math>%flooding = \frac{u_v}{u_f}</math>

If the hole pitch is unsatisfactory, return to step 6.
=====12. Check entrainment=====
''If too high, return to step 4'' Use the graph below to determine entrainment from FLV.
[[File:entrainment.jpg|400px|center|]]

The value for fractional entrainment can be used to re-estimate the column efficiency, and reevaluate the number of trays needed. Can return to step 1 for more accurate estimates.

=====13. Optimize design.=====
After returning to step 1 to reevaluate the number of trays, it is valuable to repeat steps 2 through 12 to find the smallest diameter and plate spacing acceptable at the lowest cost.

=====14. Finalize the design.=====
Optional: draw up the plate specification and sketch the layout of the plate.

====Bubble-Cap Trays====
Bubble-cap trays consist of a weir around each hole in the tray which is covered with a cap that has holes or slots to allow vapor passage. Entrainment is about three times larger than a sieve tray. Bubble-cap trays require larger tray spacing than sieve tray design. Bubble-cap trays have been known to have problems with coking, polymer formation, or high fouling mixtures. Recently, very few new bubble-cap columns are being built due to the expense and marginal benefits. However, engineers will likely encounter bubble-cap columns still currently in operation.

====Flow Patterns====
Cross flow columns are the most common pattern for distillation columns. For liquid flows between 50 and 500 Gal/min, a cross flow column is appropriate. When liquid flow is increased above 500 Gal/min, an engineer should consider designing a double pass or multi-pass column. This will reduce the liquid gradient on the tray and reduce the downcomer loading (Wankat, 2012).

===Column Sizing===
Column height will be dependent on the amount of trays required and the spacing between the trays. Normally, tray spacing of 0.15 m to 1 m is used. For columns, above 1 meter in diameter, 0.5 m can be used as an initial estimate.

Column diameter is influenced by the vapor flow rate in the column. The trays can not have excess liquid entrainment or high pressure drops; therefore, vapor velocity in the column must be maintained at a reasonable level.

An equation based on the Souders and Brown equation can be used as an estimate for the max allowable superficial vapor velocity,

<math>\hat u_v = (-0.171l_t^2 + 0.27l_t - 0.047){\frac{\rho_L - \rho_v}{\rho_v}}^{1/2}</math>

where <math>l_t</math> is the plate spacing in meters, <math>\rho_L</math> is the density of the liquid stream, and <math>\rho_V</math> is the density of the vapor stream.

Column diameter, <math>D_c</math>, can then be estimated using the relation,

<math>D_c = \sqrt{\frac{4\hat{V_w}}{\pi\rho_v\hat{u_v}}}</math>

where <math>\hat{V_w}</math> is the maximum vapor rate in kg/s (Towler et al., 2013).

===Distillation Applications===

Distillation is a process that can be implemented in various scales. There is both laboratory scaled distillation as well as very large industrial distillation. Other applications for distillation include food/alcohol processing and herb distillation for the perfume and medical industries. Typically laboratory scaled distillation occurs in batches whereas industrial distillation (e.g. fractional distillation of crude oil) occurs continuous with a constant distillate and bottom effluent streams.

Some applications of distillation are concerned the top stream only, some the bottom stream only and others both streams can be used for future products. In alcohol distillation for example, the water that is separated from the ethanol/water binary solution is discarded as waste water. In fractional distillation of crude oils, the heavy hydrocarbons at the bottom of the column are collected and sold along with the light hydrocarbons that appear in higher side draws (Wankat, 2012).

===Example Case: Ideal Distillation===

Assume an equimolar mixture flowing at 10 mol/s of 20 mol% n-pentane, 30 mol% n-hexane, and 50 mol% n-heptane. Separate the mixture into 3 products: 99% pure n-pentane, 99% pure n-hexane, 99% n-heptane. Assume the feed and products are all liquids at the bubble points. There are two process alternatives to consider in this example. The direct sequence removes the most volatile species, pentane, in the first column, and then separates hexane and heptane in the second column. The indirect sequence separates the heaviest product, heptane, and then separates pentane from hexane in the second column. This example will consider the direct sequence. Next, we must decide if these species exhibit fairly ideal behavior during distillation. Since the n-alkanes have very similar properties, it is safe to assume they will display close to ideal behavior. The next step is to look up the boiling points of the 3 species. In this case, the normal boiling points of pentane, hexane, and heptane are 309 K, 342 K, and 372 K, respectively. Also, it is a good idea to look up relative volatilites, to further verify near-ideality of the mixture, but also to obtain the information necessary for the Underwood method, which we will employ to obtain a solution. The next step is to write out material balances based on molar flows and the design specifications. They go as follows:

<math>\mu_I(nC5) + \mu_{II}(nC5) = 2 mol/s</math>

<math>\mu_I(nC6) + \mu_{II}(nC6) + \mu_{III}(nC6) = 3 mol/s</math>

<math>\mu_{II}(nC7) + \mu_{III}(nC7) = 5 mol/s</math>

<math>\mu_I(nC5) = 99\mu_I(nC6)</math>

<math>\mu_{II}(nC5) = (5/990)\mu_{II}(nC6)</math>

<math>\mu_{II}(nC7) = (5/990)\mu_{II}(nC6)</math>

<math>\mu_{III}(nC7) = 99\mu_{III}(nC7)</math>

where <math>\mu</math> represents the molar flow, and the subscript represents the product stream.

Solving this system of equations:

<math>\mu_I(nC5) = 1.985\ mol/s</math>

<math>\mu_{II}(nC5) = 0.015\ mol/s</math>

<math>\mu_I(nC6) = 0.020\ mol/s</math>

<math>\mu_{II}(nC6) = 2.930\ mol/s</math>

<math>\mu_{III}(nC6) = 0.050\ mol/s</math>

<math>\mu_{II}(nC7) = 0.015\ mol/s</math>

<math>\mu_{III}(nC7) = 4.985\ mol/s</math>

At this point we have enough information to use Underwood's method to estimate the minimum vapor flows in the column. The following three equations are used in Underwood's method:

<math>\sum_i \frac{\alpha_{ik}}{\alpha_{ik}-\phi}f_i = (1-q)F</math>

<math>(R_{min}+1)D = \sum_i \frac{\alpha_{ik}}{\alpha_{ik}-\phi}d_i = V_{min}</math>

<math>\bar R_{min}B = -\sum_i \frac{\alpha_{ik}}{\alpha_{ik}-\phi}b_i = \bar V_{min}</math>

where <math>\alpha_{ik}</math> is the relative volatility of species <math>i</math> to species <math>k</math>, <math>f_i</math> the molar flow of species <math>i</math> in the feed, <math>q</math> the fraction of the feed that joins the liquid stream at the feed tray, <math>F</math> the total molar flow of the feed, <math>D</math> the molar flow of the distillate, <math>R_{min}</math> the minimum reflux ratio <math>(=L_{min}/D)</math>, <math>d_i</math> the molar flow of species <math>i</math> in the distillate, <math>V_{min}</math> the minimum vapor flow possible in the top section of the column to accomplish the desired separation, <math>\bar R_{min}</math> the minimum reboil ratio <math>(=\bar V_{min}/B)</math>, <math>b_i</math> the molar flow of species <math>i</math> in the bottoms product, and <math>\bar V_{min}</math> the minimum vapor flow in the bottom section of the column. The final variable, <math>\phi</math>, will be solved for using the first Underwood equation, and it's value will be decided based on the relative volatilities of the key components in the column.

So, after solving the first Underwood equation, we get two values for <math>\phi</math>, 3.806 and 1.462. Because 3.806 is between the relative volatilities of the key components, we will substitute that value for <math>\phi</math> into the second Underwood equation. Doing so for both columns gives <math>V_{min} = 6.4\ mol/s</math> for the first column and <math>V_{min} = 8.9\ mol/s</math> for the second column, for a total minimum vapor flow of 15.3 mol/s. The process would then be repeated for the indirect sequence, and the decision for which process to use would be justified by the process with the overall minimum vapor flow (Biegler et al., 1997).

==Absorption==
===Description of Absorption===
Another separation process used in industry is absorption, which is used to remove a solute from a gas stream. It accomplishes this by contacting the gas mixture with a liquid solvent that readily absorbs the undesirable components from the gas stream, purifying the gas stream. This separation process is determined by the inputs of the liquid flow rate, temperature, and pressure.

The absorption factor, which can be determined mathematically, determines how readily a component will absorb in the liquid phase. The absorption factor of component i is

<math>A_i=L/K_iV</math>

where <math>L</math> is the liquid flow rate entering the column, <math>V</math> is the vapor flow rate entering the column, and <math>K_i</math> is the vapor/liquid equilibrium ratio for component i (Peters & Timmerhaus, 2003). Higher absorption factors result in higher absorptivity into the liquid and a decrease in the number of trays required for separation, however a diminishing return occurs after the absorption factor is greater than 2.0. An absorption factor of 1.4 is most commonly used.

In general absorption can be seperated into two overarching categories, physical and chemical absorption. In physical absorption, the unwanted solute in the gas is absorbed into the liquid phase because solubility of the component is higher in the liquid phase than the gas phase. In chemical absorption the solute is removed from the gas via a reaction with the solvent, this reacted product is then transported into the liquid phase (Danckwerts 1965). There are two types of chemical absorption reversible and irreversible. Generally reversible chemical absorption is preferred as the solvent can be put through a stripper and regenerated so it can be recycled back to the absorption process (Wankat, 2012).

===Absorption Apparatus===

There are five major apparatus used for absorption in industrial application. These five pieces of equipment are spray absorbers (or towers), ejector (venturi) scrubbers, packed columns, trayed columns, and film absorbers (Schmidt, 2012).

==== Spray Tower vs Ejector Scrubber ====

In both '''spray tower''' and the '''ejector scrubber''' nozzles are employed to produce small solvent droplets. These small droplets increase the surface area of the liquid to gas contact allowing for the maximum amount of mass transfer to occur between the gas mixture and the liquid. The major difference between the two nozzle equipment designs is the configuration and type of nozzles. In the ejector scrubber shown in Figure 3 there is a single nozzle that is generally a higher pressure spray nozzle that produces finer solvent drops allowing for an even greater amount of mass transfer enabling better physical absorption (Schmidt, 2012).
[[File:Ejectorventuri.jpg|thumb|200px|center|Figure 3. Ejector Scrubber (US EPA, 2006)]]
'''Spray towers''' on the other hand generally have many nozzle at different heights where the liquid solvent will be sprayed out of to contact the gas running through the tower. This design is used in order to ensure the gas contacts the liquid as throughout the tower. These nozzles are lower pressure than a ejector scrubbers nozzle and thus physical mixing is worse in this configuration. Since physical mixing is generally worse in this configuration it is usually used in conjunction with a chemical absorption process. The other major difference between the ejector scrubber and the spray tower is that gas and liquid flow is cocurrent in the former while it is countercurrent in a spray tower. A spray tower absorber is shown below in Figure 4 (Schmidt, 2012).
[[File:SparyTowerAbsorber.jpg|thumb|200px|center|Figure 4. Spray Tower Absorber (US EPA, 2006)]]

==== Tower Type Absorption Apparatus ====
'''Packed column absorbers''' and '''tray column absorbers''' have very high efficiencies for the removal of an unwanted solute in the gas stream. The major disadvantage a trayed column has when compared to a packed column is the pressure drop. The pressure drop in a packed column is generally very low, whereas in between each tray of a trayed column pressure drop can be quite large. However the advantages inherent to trayed columns become clear when one needs the solvent to have a high concentration of the component to be removed from the gas stream. This is most important in the case where there is a very low concentration of the component in the gas stream and the specification states the solvent must contain a high concentration of that component. In this case the flow rate of the solvent may not be high enough for a packed column, however in a trayed column the solvent flow rate can be near zero for operation (Schmidt, 2012). Packed and trayed column internals are very similar to the setups found in the respective distillation columns.

For a '''trayed column''' the plate efficiency can be calculated using O'Connell's Correlation which invovles the Henry's Law constant, total system pressure, and solvent viscosity at the operating temperature (Towler & Sinnott, 2013).

<math>x=0.062*\frac{\rho_s*P}{\mu_s*H*M_s}</math>

where
<math>x</math> is the tray efficiency,
<math>\rho_s</math> is the density of the solvent in <math>kg/m^3</math>,
<math>P</math> is the total pressure of the system in <math>N/m^2</math>,
<math>\mu_s</math> is the solvent's viscosity in <math>mNs/m^2</math>,
<math>H</math> is the Henry Law constant in <math>1/(Nm^2*(mol fraction))</math>,
and <math>M_s</math> is the molecular weight of the solvent.

A packed towers height can be determined using the equations below when concentration of solute is below 10% so that the assumption that the flow of gas and liquid will be essentially constant throughout the column holds (Towler & Sinnott, 2013). The height of packing <math>Z</math> is given by the following equation:

<math>Z=\frac{L_m}{K_G*a*P}*\int\limits_{y_2}^{y_1} \frac{dy}{y-y_e}\,</math>

where <math>P</math> is the total pressure, <math>a</math> is the interfacial surface area per unit volume, <math>y_1</math> and <math>y_2</math> are the mol fractions of the solute in the gas stream at the bottom and top of the column respectively, <math>G_m</math> is the molar gas flow rate per unit cross-sectional area, and <math>y_e</math> is the mole fraction of solute in the gas that would be in equilibrium with the liquid concentration.

The first half of the equation before the integral can be called the height of an overall gas-phase transfer unit <math>H_G</math> and the second part of the equation is the number of overall gas-phase transfer units or <math>N_G</math>. Using these definitions the above equation can be simplified to

<math>Z=H_G*N_G</math>

These equations assist in sizing an absorption column (Towler & Sinnott, 2013).

==== Film Absorber ====
The final absorber the film absorber is generally used in the case where the heat of absorption must be removed. The film absorber operates by sending the gas and solvent through a heat exchanger where the solvent creates a thin film on the walls of the tubes and the gas flows through the interior allowing for solute transfer. The good heat transfer present in a film absorber makes it preferable for situations where low temperatures are required for a high recovery of the solute (Schmidt 2012).

===Industrial Absorption Processes===
An industrial example is lean oil absorption, which is used to separate nitrogen and other impurities from natural gas. A lean oil is contacted with low quality natural gas, and the methane is selectively absorbed by the lean oil, leaving the impurities behind. The methane is subsequently regenerated from the rich oil as high quality natural gas (Petrogas Systems, 2014).

Other common industrial practices of absorption come from sour gas treatment. Amine gas treating is used to remove hydrogen sulfide or carbon dioxide from gas streams via a reversible chemical absorption. In amine gas treating the sour gas is fed to the bottom an absorber where amine solution is fed to the top along with any necessary make up water. The sour gas components are absorbed into the amine via a chemical absorption method. Sweet gas leaves the top of the absorber whereas the amine out of the bottom, now rich with acidic components is sent to a regenerator where the acid gas components are stripped and the acid gas is generally sent to a flare whereas the amine now lean again is recycled back into the first absorber (Miller & Zawacki, 1978). Figure 5 below shows the typical setup of an amine plant. Another type of sour gas treatment that uses absorption is Merichems LO-CAT process which uses a chelated iron to remove hydrogen sulfide from feed gas in the absorption column (Merichem 2015).
[[File:AmineTreating.png|thumb|400px|center|Figure 5. Amine Gas Treating Plant Schematic]]

==Stripping==
This process separates solutes from solvents (often after absorption, to purify the solvent so that it can be recycled to an absorber). Stripping will depend on the vapor and liquid flow rates, as well as the temperature and pressure of the column. There is a temperature drop down the column, so columns generally have either an increased operating temperature or decreased operating pressure.

The stripping factor of component i is

<math>S_i=K_iV/L</math>

where <math>K_i</math> is the vapor/liquid equilibrium ratio, <math>V</math> is the vapor flow rate entering the column, and <math>L</math> is the liquid flow rate entering the column, will determine how much of solute i will be stripped from the liquid into the vapor phase (Peters & Timmerhaus, 2003). The usual range for the stripping factor is between 1.2 and 2.0, with a stripping factor of 1.4 being most economic.

An example of stripping in industry is the deodorization of food items such as oils. The oil is heated and allowed to trickle down the column while steam flows up from the bottom of the column. At the vapor-liquid interface, volatile components of the oil transfer to the steam and are carried off the top of the column, leaving a purified oil product (Alfa Laval, 2014).

==Bioseparations==
===Importance===
As our ability to manipulate and engineer biological systems improves, biological products are becoming an increasingly important source of therapeutics and fuels. The production of fuels from biomass via either the enzymatic breakdown of a feedstock or the secretion of usable lipids from algae is a promising new energy source. Additionally, enzymes, antibodies and other therapeutic proteins have been applied to the treatment of a wide range of diseases. Although each process requires its own set of separations, all follow the same basic format: separation of biomass, product isolation, and product purification (Belter et al., 1998). This section will provide examples of unit operations in each step. Ultimately, the choice of separation process and unit operations will depend on the specific process and product. The descriptions below are examples of the most common bioseparation operations within the general platform (Harrison et al., 2003).

Bioprocesses begin with fermentations or growth operations. In biofuel production processes, this may involve growing algae or breaking down corn or cellulosic biomass. For the production of therapeutics, mammalian or bacterial cells may be grown in a fermentor and the product secreted into the supernatant or harvested from the cells.

===Biomass Separations===
After fermentation and product production, the solid biomass must first be separated from the desired product. If the product is secreted from the cells, this can be done immediately after fermentation ends. If the product is not secreted, the cells must first be lysed.
Cell lysis is the process of lysing, or breaking, the cell in open. Mechanical lysis is the simplest, and involves physically breaking the cell either by mashing (think mortar and pestle) or blending the cells into a homogenous solution in a homogenizer. Chemical lysis is another method, achieved by introducing an osmotic shock or chemically degrading the cell membrane. Additional separation can be achieved by flocculation, which is the process of aggregating biomaterial by charge neutralization or bridging. These larger complexes are easier to separate from smaller molecules (Harrison et al., 2003).

The next step is removing the unwanted biomass from the product in solution. Separation by centrifugation or sedimentation are the most common, although filtration is sometimes also used for processes where a biomass cake is desired. Both methods utilize density differences to separate the product from the solid biomass (Towler and Sinnott, 2013).

====Sedimentation====

Sedimentation relies purely on the force of gravity, while centrifugation speeds the settling process by subjecting the cells to a centrifugal force. Sedimentation in a settling tank is the simplest method of solid-liquid bioseparation. In this process, biomass in a tank is simply allowed to settle to the bottom over time. While this process is inexpensive, requires little energy and can separate out large volumes of biomass, it generally requires long time periods and is only mostly in very large-scale processes where active centrifugation is difficult (Belter et al., 1998).

====Centrifugation====
Centrifuges are widely utilized across many processes, and thus a wide variety of scales and designs have been developed. Disk-stack centrifuges, in which the solid phase is deposited onto “shelves” in the center of the spinner and liquid phase is pushed to the outside, are some of the most commonly used centrifuges in industry. They are especially suited to biomass separation processes because they can be built on a large scale and are ideal for separating fine solids from liquids. [[File: Disk_stack_centrifuge_towler.png|frame|center|Fig. 6: Diagram of a disk-stack centrifuge (Tolwer et al, 1997).]] Tubular bowl centrifuges are also common and can reach separation efficiencies of up to 90%. Heavier products accumulate along the sides of the bowl, while the light phase flows out the top. They separate products by can be used both to separate solids from liquids and immiscible liquids, such as and oil product and an aqueous broth (Tolwer and Sinnott, 2013). [[File: tubular bowl centrifuge towler.png|frame|center|Fig. 7: Diagram of a tubular bowl centrifuge centrifuge (Tolwer and Sinnott, 2013).]]

Centrifugation scale-up is made easier by sigma analysis, which allows for the estimation of appropriate feed rates for different size centrifuges. The sigma factor is dependent on the inner and outer radius of the centrifuge, the angular velocity, and the sedimentation velocity of the solid particles being separated. It can be thought of as the characteristic cross-sectional area with units of [length]2. The sedimentation velocity can be calculated by

<math>v_g={\frac{2a^2(\rho-\rho_0)}{9\mu}}g</math>

where <math>v_g</math> is the sedimentation velocity, <math>a</math> is the cell or biomass particle diameter, <math>\rho</math> is the particle density, <math>\rho_0</math> is the fluid density, and <math>\mu</math> is the fluid viscosity. The volumetric flow <math>Q</math> can be estimated by

<math>Q=(v_g)(\Sigma)</math>.

The equality

<math>{\frac{\Sigma_1}{\Sigma_2}}={\frac{Q_1}{Q_2}}</math>

can be an easy way to estimate equivalent flow rates between a small-scale centrifuge 1 and larger centrifuge 2 (Harrison et al., 2003).

====Example: Centrifugation Scale-up====

You are trying to separate a cell of radius 0.4 <math>\mu</math>m with a density of 1.05 g/cm3 from broth of mostly water (density of 1 g/cm3 and viscosity of 0.01 g/cm s). The sigma factor of the centrifuge you are using is 1 x 106 cm2. A] What volumetric flow rate should you use? B] If you want to scale up the process to a centrifuge with <math>\Sigma</math> = 3 x 106 cm2, what flow rate would you use in the larger centrifuge?

Solution:
A] Using the equation for <math>v_g</math>, and being mindful of units, the sedimentation velocity equals 1.74 x 106 cm/s. The flow rate, then, equals

<math>Q=(1.74 x 10^-6)(1,000,000) = 1.74 cm^3/s = 0.104 L/min</math>.

B] Keeping in mind that for the same process, <math>v_g1 = v_g2,</math> and rearranging the sigma factor equality, the new flow rate is

<math>Q_2 = {\frac{\Sigma_2 x Q_1}{\Sigma_1}} = {\frac{(3 x 10^6)(0.104)}{1 x 10^6}} = 0.313 L/min </math>

===Product Isolation===
Liquid-liquid separation, to extract the product from the aqueous phase, is much less straightforward than liquid-solid extraction. Many methods - especially adsorption, filtration, and precipitation - are similar in principle to operations found in other, non-biological separations. The exact separations used depend on the nature of the product and the scale of the process. These processes are nearly identical to their non-biological counterparts, and their description is left to other sections.

Particular care needs to be taken with protein products because of their instability, and the selection of an appropriate solvent or adsorbent is crucial to a successful process (Harrison et al., 2003).

===Product Purification===
The final steps of protein purification and polishing remove any remaining contaminants and bring the concentration of product to an appropriate value for applications. Purification processes for food-grade and medical products can be extensive, as sterility and high purity are essential. Purification in fuel-producing processes may be less extensive, depending on the process. Chromatography and crystallization are two common steps in purification and are especially used in industrial scale protein production. Several different types of chromatography exist with the ability to carry out different types of separations.
Chromatography is similar to adsorption in that it relies on differences in affinity between solutes and a solid surface. A solution is eluted through a column containing a solid resin with various affinities for the substances in solution. In adsorption, the solutes are evenly saturated throughout the column. Chromatography differs in that solutes are deposited a resin phase before the column is flushed with an elution solvent specific that results in solutes eluted in bands.

==== Ion Exchange Chromatography ====

There are two main types of ion exchange columns—anion and cation. Anion exchange resins have a positive charge and are used to retain products with a negative charge. Cation exchange resins have a negative charge and are used to retain products with a positive charge. The pH of the elution buffer is change to force a specific solute to wash out, depending on whether the pH of the buffer is above or below the isoelectric point of the solute (Belter et al., 1998). This is especially useful for the separation of protein product (including antibodies), nucleic acids, and other charged molecules. When the solutes have sufficiently different isoelectric points, the pH of the buffer is manipulated to affect the solute charge and force the product to elute while the solute remains preferentially bound to the resin, or vice versa (Harrison et al., 2003). In general, the most strongly charged molecules will remain in the column for a longer period of time. Elution washes through the weakly bound ions before the more strongly bound ions. Different speeds of elution can be visualized as in figure 8.

[[File:chromatography.png|frame|center|Fig. 8: Illustration of product bands in an elution chromatography column (Belter et al., 1998).]]

==== Size Exclusion Chromatography ====

In gel filtration chromatography, small molecules are "trapped' by the porous resin and take longer to flow through the column. Larger products will elute first because the smaller molecules are better able to penetrate the resin. This forces them to take a much longer path through the column, which means it takes longer for them to elute. This operation is often used when there is a distinct difference in size between the desired product and other solutes.

==== Affinity Separations ====

Affinity chromatography is very similar to ion exchange chromatography in that the interactions between the material in the column and the molecules in the feed. The main difference is that affinity chromatography can rely on a great variety of types of interactions. Two very common types of affinity are exploited in affinity chromatography columns. The first is immunoaffinity. Proteins are specifically bound by antibodies which can be incorporated onto beads and used in chromatography. Antibodies are designed to bind only a single protein, so these interactions are considered to be highly specific. The protein can be eluted using a buffer that changes the pH or salinity in the column, which adversely affects binding (Hage, 1999).

Proteins can be separated from very complex and unrefined mixtures using this method because of the great specificity. Antibody-protein interactions are so specific that individual proteins can be isolated from blood samples as shown in the figure below. The main drawback of this method is simply that specific interactions between proteins and binding targets do not always exist. Antibodies exist for many proteins, but for others they must be customized or created using some sort of evolution process. This is not a trivial task and can require a significant capital investment (Hage, 1999).

[[File:Affinity_Chromatography_Example.PNG|frame|center|Fig. 9: Affinity purification separating fibrinogen from human plasma using an anti-fibrinogen antibody (Hage, 1999).]]

The other main type of affinity chromatography is based on protein specific tags and the molecules or surfaces to which they bind. One of the most common types of protein tags used is the polyhistidine tag. This tag consists of 6-8 consecutive histidine residues which can be added to the exterior of the desired protein product. The addition of this tag requires alterations to the coding sequence of the protein. The polyhistidine tag binds strongly with nickel and cobalt ions. The product with the tag can then be eluted with imidazole—a small molecule with the same structure as the functional group of the amino acid histidine. Imidazole will bind the cobalt and nickel ions more strongly than the histidine in the tag. Along with chromatography, protein tag interactions can be leveraged with the use of beads that can be deposited directly into the solution containing the protein of interest (Lichty et al., 2005).

Different proteins can be purified in this manner with varying levels of efficiency. There is a very high dependence on the size of the protein, electronic properties, and steric considerations. The figure below demonstrates varying separation efficiencies with proteins of different sizes (HisPur Cobalt Resin - Thermo Fisher Scientific, 2016).

[[File:HisTagPurification.PNG|frame|center|Fig. 10: Protein gel demonstrating the separation of proteins from a sample. The far left column represents the most efficient separating conditions using cobalt and the far right column represents a negative control with no purification performed (HisPur Cobalt Resin - Thermo Fisher Scientific, 2016).]]

Several other types of tag-bead interactions can be utilized in separations processes. Maltose Binding Protein is a small protein that can be added to a protein of interest. It binds strongly with beads coated in immobilized maltose and can be released by flushing with maltose. As MBP is a full sized protein that typically must be removed from the protein of interest in order for it to be used. In this case, the site specific TEV protease is often used cleave MBP from the protein of interest. In addition, under specific circumstances, other unique tags can be used and provide varying levels of specificity in separations. The Flag tag, 3x Flag tag, Glut tag, and Strep tag. While these are all commonly used, the polyhistidine tag is the most popular because it gives the highest level of specificity. Some of the main purification tags are compared in the table below (Lichty et al., 2005).

{| class="wikitable"
|-
! Tag
! Size (aa)
! Resin
! Eluting Agent
! Capacity (mg/ml)
! Cost/10 mg
|-
| His
| 6
| Talon
| Imidazole
| 5-14
| $18
|-
| Maltose Binding Protein
| 396
| Amylose
| Maltose
| 3
| $12
|-
| Glutathione S-Transferase
| 218
| GSH-Sepharose
| Glutathione
| 10
| $36
|-
| Strep II
| 8
| Strep-Tactin Sepharose
| Desthiobiotin
| 0.1
| $293
|-
| FLAG
| 8
| Anti-FLAG M2 MAb Agarose
| Flag peptide
| 0.6
| $1045
|}

Unfortunately, based on the contents of the solution, the characteristics of the protein, and the specific product being used, there are extremely large amounts of variation in yield and purity for each type of tag. Therefore, it is crucial to allocate time towards preliminary tests and scale-up experiments with the exact solution from which the protein of interest is being separated. The his tag is typically the most used because, as seen in the table above, it has one of the lowest costs and highest yields. In addition, because it is only a 6 amino acids, it often is not cleaved from the protein of interest which minimizes the number of steps required in the overall purification process (Lichty et al., 2005).

==== Crystallization ====

Crystallization, or the formation of solute crystals from a solution, is especially useful in biomolecule separations because it is possible to obtain a 99.9%+ product purity. In crystallization, a diluent is added to the homogeneous solution that reduces the solubility of the product to the point that it “falls out” of solution and crystallizes. It is similar to precipitation but results in the formation of crystals rather than unordered aggregates. Crystallization can be used on a laboratory scale for determining protein structure, on on the industrial scale for antibody and therapeutic protein productions. Batch crystallizers are often used in industry because of their simplicity and inexpensiveness compared to continuous crystallization (Harrison et al., 2003).

==Membrane Separation==
Membrane separation takes advantage of the selective permeability of membranes; they allow certain particles to pass through and selectively stop other, generally unwanted, particles. The component that passes through is called the permeate and the component stream that is rejected is called the retentate or concentrate. The applicability of membranes comes from the fact that their selectivity is determined by their pore size, which can be controlled during the creation of the membranes. Additionally, Membrane processes do not require heat meaning they generally require less energy than conventional separations technology such as distillation and crystallization. Membrane separations processes are generally classified as microfiltration, ultrafiltration, or nanofiltration depending on the size of the particles to be filtered out.

[[File:membrane techs.png|frame|Figure. Cutoffs for different membrane categories]]

===Membrane Selection, Construction, and Flow Geometries===
Membrane permeability and selectivity are the two most important factors to consider when selecting a membrane. For gas separations, the permeation of the gas is usually facilitated by the gas dissolving in the membrane on one side and then evaporating on the permeate side. Therefore permeability depend largely on the solubility of components in the membrane.

The two most commonly utilized membrane configurations are hollow fiber and spiral wound. Hollow fiber is generally the most commonly utilized module for gas separations. These are formed by gluing the two ends of the hollow fiber to a resin forming a closure. The fibers are housed in a shell much like a heat exchanger. The feed flows past thousands of tubes with the permeate flowing into the hollow tubes and out the closure. The retentate then flows out of the shell not having gotten through the membrane.

[[File:hollow membrane.jpg|frame|Figure. Hollow fiber membrane module]]

Spiral wound membranes are created by sealing two membrane sheets back to back on three edges to form a sort of pocket. This fourth open edge is then attached to a porous tube which allows permeate to go through it. Several membrane pockets are attached to a single tube and wrapped around in a spiral.

[[File:spiral membrane.jpg|frame|Figure. Spiral wound membrane module]]

Flow geometry is usually either dead-end geometry or cross flow geometry. In dead end, the fluid flow is normal to the membrane surface while cross flow is parallel to the membrane surface. Dead end geometry is usually used with hollow fiber membranes while cross flow is used with spiral wound membranes. Each geometry has advantages and disadvantages. Dead end geometry is generally cheaper to set up and therefore has lower initial capital costs. However, it is very vulnerable to membrane fouling, which reduces the effectiveness of the membrane. This is usually the geometry set up for small scale lab experiments. The tangential flow devices are more cost and labor-intensive, but they are less susceptible to fouling due to the sweeping effects and high shear rates of the passing flow. Most commercial industrial membrane separations are done using spiral wound cross flow membrane modules.

===Applications===
====Food Industry====
Due to the fact that MD can be conducted at relatively low feed temperatures, it was successfully tested in many areas where high temperature applications lead to degradation of the process fluids especially in food processing. It was demonstrated that MD can be used for the concentration of milk, for the recovery of volatile aroma compounds from black currant juice, and for the concentration of many other types of juices including orange juice, mandarin juice, apple juice, sugarcane juice, etc.
====Reverse Osmosis====
Reverse osmosis is the most widely used membrane separation process. In this process, fresh water passes through the membrane while dissolved salts and other solids are rejected and stay in the concentrate. In this process, feed water is pressurized in order to overcome the osmotic potential difference between the salty retentate and the fresh water desired. These processes are generally run using spiral wound membrane cylinders using a cross flow setup.

===Membrane Model===
The two most important components when considering different membranes are the permeability, which will determine flux through the membrane, and selectivity, which will determine what passes through the membrane and how much. The flux through a membrane is defined as:
<math> M_i = \frac{P_i}{δ}(p_{i,f} - p_{i,p})</math>
Where Mi is the molar flux of component i, Pi is the permeability of the membrane for component i, δ is the membrane thickness, and pi,f and pi,p are the partial pressures of component i on the feed side and permeate side respectively.
The average flux across a long cylindrical membrane such as the spiral wound module is given by:
<math> \int_0^Lm \frac{M_i,dx}{L_m}</math>
Where Lm is the length of the cylinder and x is length in meters

Membrane selectivity of the ideal separation factor is given as the ratio of the permeability of one substance over another as shown:

<math> S_(i,j) = P_i/P_j </math>

where Sij is the selectivity of the membrane for component i over j.

==Cyclones==
Centrifugal Separators, more commonly know as Cyclones, are one of the most widely used gas-solid separators. They are typically used for particles that are 5μm or larger, but if agglomeration occurs they can sometimes separate particles as small as particles as small as .5μm. These cyclones can be designed for high efficiency separations or for high throughput (Towler, 2012).

[[File:cyclone.png|thumb|center|Reverse-flow Cyclone (Towler, 2008)|500x300px]]

===Design===
Design Procedure
There are multiple design methods for cyclones. The deign method discussed below is the Stairmand method. This involved creating two designs one of which was for high throughput and the other which was designed for high efficiency. Using these base designs future cyclones can simply be estimated using scaling factors (Towler, 2012).

1. Select whether efficiency or performance is more important for your design

2. Determine the particle size distribution in your stream

3. Estimate the number of cyclones needed

3. Calculate the cyclone diameter

4. Calculate the scale up factor

5. Calculate the cyclone performance and overall efficiency

6. Calculate the cyclone pressure drop

7. Cost the system

====Particle Size====
The diameter of particles separated by cyclones are governed by this design equation. This equation is dependent upon the standards of the cyclone and how much you are trying to deviate from its standard operating conditions. Looking into the equation it shows that a larger internal cyclone diameter, higher flowrate, higher difference in density, and lower viscosity result in smaller particles being seperated. The whole term after d1 is known as a scale up factor. When designing a cyclone you would determine the flow rate, viscosity, particle size, and density difference would be known prior to using this design equation, so this would be used to optimize the diameter of the cyclone you were designing (Towler, 2012).

[[File:Towler 10.8.PNG|center|Reverse-flow Cyclone (Towler, 2012)|500x300px]]

d1=mean diameter of particle separated at the standard conditions, at the chosen
separating efficiency

d2= mean diameter of the particle separated in the proposed design, at the same
separating efficiency

Dc1 =diameter of the standard cyclone

Dc2 =diameter of proposed cyclone

Q1 =standard flow rate

Q2 =proposed flow rate

Δρ1 =solid-fluid density difference in standard conditions

Δρ2 =density difference, proposed design

μ 1=test fluid viscosity

μ 2=viscosity, proposed fluid.

[[File:Towler10.44.PNG||thumb|center|(a) Standard Cyclone Dimensions (b) High Gas Rate Cyclone (Towler, 2012)|500x300px]]

====Efficiency====
Efficiency is the percentage of total solids that are removed from the stream The graphs below were experimentally for the two standard designs Stairmand Method. This graph has the efficiency vs the particle size.

[[File:Towler10.45.PNG||thumb|center|Performance curves, standard conditions. (a) High-efficiency cyclone performance curves, standard conditions. (b) High gas-rate cyclone (Towler, 2008)|600x400px]]

If you have to scale the efficiency it can be estimated by the point where a horizontal line from the efficiency of the original particle size intersects the new particle size. A demonstration is on the graph below.

[[File:Towler 10.46.PNG||thumb|center|Scaled Performance Curve(Towler, 2008)|500x300px]]

====Pressure Drop====
One of the most important design factors in a cyclone is its pressure drop. The pressure drop occurs because entry and exit losses, kinetic energy losses, and friction losses that occur inside the cyclone. The pressure drop can be calculated using the following equation.

[[File:Towler 10.9.PNG|center|Reverse-flow Cyclone (Towler, 2008)|500x300px]]

The variables are defined as follows:
DP = cyclone pressure drop, millibars;
rf = gas density, kg/m3
;
u1 = inlet duct velocity

u2 = exit duct velocity

rt = radius of circle to which the center line of the inlet is tangential

re = radius of exit pipe

φ =factor from figure below

ψ = fc(As/A1)

fc = friction factor

As = surface area of cyclone exposed to the spinning fluid

A1 = area of inlet duct

[[File:Towler_10.47.PNG|thumb|center|Cyclone Pressure Drop Factor (Towler, 2008)|600x400px]]

====Cost Estimation====
The following values are represented in 2002 dollars. Cost is mostly a function of the amount of gas that need to be passed through the cyclone and the units are in standard meters cubed per second (sm3/sec).

The costs can be broken down into three categories:

Capital Costs: $4600-$7600 sm3/sec

Operations and Maintenance: $1500-$18,000 sm3/sec

Annualized Cost: $2800-$29000 sm3/sec

This all comes to $.47-$440 per metric ton of pollutant removed. The high variation in these costs have to do with smaller units being much less efficient per unit of pollutant removed. Furthermore the cost is also effected by the amount of particulates that are in the gas to begin with. For example a gas with 10 g/sm3 and air with 100sm g/sm3 would have similar pumping costs, but very different cost per unit of solids removed (EPA).

==Other Separation Processes==
===Extraction===
Liquid-liquid extraction is a process for components with overlapping boiling points and azeotropes. The process requires a solvent such that some of the components of the mixture are soluble, and then the components will be separated based on this solubility in the liquid. This process can operate at moderate temperatures and pressures, so is not very energy intensive. However, a distillation column is required to extract the solvent for recycle. More recently, supercritical fluids have replaced liquid solvents in some processes for L/L extraction, due to the solute’s ability to more rapidly diffuse through them. The issue with these fluids, however, is that they must be operated at extremely high pressures and temperatures, increasing both capital and operating expenses of the process (Peters & Timmerhaus, 2003).

===Crystallization===
This process recovers solutes that have been dissolved in solution. The resulting product is in the solid phase. Depending on the material properties of the solute and solvent, the solute is recovered by precipitation after cooling, removal of solvent, or adding precipitating agents. Crystallizers are designed based on phase equilibria, solubilities, rates and amounts of nuclei generated, and rates of crystal growth. Every crystallization process is a unique system, so plant evaluation is usually required before complete implementation. Crystallization can be performed in both batch and continuous processes, and design features can control crystal size to an extent (Peters & Timmerhaus, 2003).

===Membrane Separation===
This separations process uses selectively permeable membranes to separate components in a mixture. Typically, one of the components will freely pass through the barrier while the other components will not. The stream that passes through the membrane is the permeate and the stream that does not pass is the retentate. The driving force behind this separation is a pressure gradient. Membrane separation is beneficial because it can separate mixtures at the molecular and small particle level. Furthermore, there is no phase change required so the energy input is low. Limitations of this process include achieving high product purity, incompatibility with certain stream components, low operating temperature, and low flow rates. Although membrane separation is generally not scaled up, examples of scaled-up membrane separation include seawater desalination and hydrogen recovery (Peters & Timmerhaus, 2003).

===Adsorption===
Adsorption involves an adsorbent and adsorbate. The adsorbent is typically a solid, and will typically separate the adsorbate from the stream. This process usually includes a desorption step that regenerates the adsorbent for further use. Raising the temperature or increasing the concentration of the adsorbate can reverse the adsorption process. Although the recycle of the adsorbent is a very economic design feature, the downside of this step is that it results in a cyclic process, which introduces complexity to the overall process. Industrial applications of this process are for bulk separations and gas purification. The adsorption/desorption process in these situations involves a large amount of heat transfer, which design engineers must take into account when sizing and selecting equipment material (Peters & Timmerhaus, 2003).

===External Field/Gradient Separation===
These separations use external force fields or temperature gradients to separate responsive molecules or ions. The use of these processes is fairly limited to a few specialized industrial applications (Peters & Timmerhaus, 2003).

===Settling and Sedimentation===
In settling processes, solid particles or liquid drops are separated from a stream by gravity. The stream can be in either the liquid or gas phase. For vapor-liquid mixtures, flash drums are generally used to separate the mixture. The velocity of the vapor must be less than the settling velocity of the liquid drops for this separation to occur. For liquid-liquid separation, the horizontal velocity of the fluid must be low enough to allow the low-density droplets to rise to the interface and the high-density droplets to move away from the interface and coalesce. In sedimentation, the result of the process is a more concentrated slurry. Typically a flocculating agent is used to aid in the settling process. One way to perform this separation is to use a cone-shaped tank with a slowly revolving rake that scrapes and moves the thickened slurry to the center of the cone for removal (Peters & Timmerhaus, 2003).

[[File:Example.jpg]]

====Clarifiers====
[[File:Circular_Clarifier.png|300px|thumb|bottom|Figure 9: Circular clarifier with some components labelled.]] [[File:Rectangular_Clarifier.png|300px|thumb|bottom|Figure 10: Rectangular clarifier with some components labelled.]]

Clarifiers are one of the methods used for the continuous removal of particulate solids from liquids through sedimentation by gravity. Applications include process water pretreatment, waste water treatment, and drinking water purification. Historically, clarifiers were originally developed to limit nutrient input into surface water due to fear of eutrophication. Today, they have a number of uses, particularly in wastewater treatment processes, metal removal, disinfection, and membrane pretreatment. The process helps removed dissolved solids, silt, and undesirable metals from the water, making it more suitable for downstream processes as well as human consumption (Wilson, 2005).

Clarifiers are typically used in conjunction with coagulation or flocculation agents, which promote dissolved particles to join into clumps and settle out of solution (Towler and Sinnot, 2012). Clarifiers typically consist of a large circular tank with a rotating rake at the base which scrapes settled solids towards the center. In the case of a rectangular clarifier, they are scraped to one side. Diagrams of both are represented in figures 9 and 10, respectively (NMED Surface Water Quality Bureau, 2015). Separated solids are allowed to settle to the bottom of the tank as a sludge, whereupon they are collected by the rake and disposed of properly. In the case of floating contaminants, it is possible for the clarifier to include a skimmer as well.

Clarifier efficiency varies with certain factors, including the settling characteristics of solids removed and the surface overflow rate of the tank. Clarifier efficiency can be found using the following relation:

<math>\begin{align}
E_{TSS} &= E_{TSSmax}\left ( 1 - e^\frac{\lambda}{SOR} \right )
\end{align}</math>

where <math>E_{TSS}</math> is the efficiency of total suspended solids (TSS) removal, <math>E_{TSSmax}</math> is the maximum possible efficiency, <math>\lambda \left [\frac{m}{d} \right ]</math> is the settling constant, and <math>SOR \left [\frac{m^3}{m^2 d} \right ]</math> is the surface overflow rate. The effect of flocculation chemicals on TSS can be seen in figure 11. However, it should be noted that chemical addition will increase sludge quantity and may have an adverse effect on plant aesthetics, which increases maintenance costs (Wilson, 2005).

[[File:Chem_Addition.png|200px|thumb|bottom|Figure 11: The effect of flocculating agents on total suspended solids removal in clarifiers.]]

=====Lamella Clarifiers=====

Lamella clarifiers use inclined plates in order to maximize the settling area for solids. Solids continue to settle into a hopper at the bottom of the tank while clarified water exits up through the inclined plates. This allows for the design of a smaller tank, which leads to large savings in capital costs. A lamella clarifier is pictured in figure 12.

[[File:Lamella_Clarifier.png|300px|thumb|bottom|Figure 12: A lamella clarifier with components labeled.]]

Typically, inclined plates are installed at an angle of 45 to 60 degrees and spaced 40 to 120 mm apart, which increases effective settling surface area by a factor of 6 to 12 compared to traditional clarifiers. For effective use, it is recommended that the Reynolds number be below 2000, Froude number higher than 10-5,and detention time be longer than 3 to 5 minutes. For this implementation, the equations are as follows:

<math>\begin{align}
N_{Re} &= \frac{VR}{\nu} \\
N_{Fr} &= \frac{V^2}{Rg}
\end{align}</math>

where <math>R</math> refers to the hydraulic radius, which is the cross-sectional area of the lamella, <math>V</math> is the liquid velocity, <math>\nu</math> is the kinematic viscosity, and <math>g</math> is the gravitational constant (Wilson, 2005).

=====Advantages=====

Clarifiers offer a proven, relatively inexpensive solution for solids removal. The chemical coagulants used are cheap and provide a low operating cost as well as simple maintenance. Construction is typically simple, leading to low capital costs and equipment that is easy to accommodate and maintain. Their design is also flexible, with various options such as skimmers and scrapers offering increased removal efficiency (Wilson, 2005). Operation of clarifier tanks also has lower energy requirements than membrane filtration for solids removal, given that most of the separation is aided by gravity. Water exiting clarifier units has a silt density index (SDI) averaging 4.0, which is low enough for further membrane treatment such as reverse osmosis (Prihasto, 2009).

=====Disadvantages=====

Clarifiers necessitate low turbulence to prevent resuspension of solids. This essentially requires a low entrance velocity, which can limit the production rate of certain processes or call for more clarifier units, which would drive up costs. Furthermore, clarifiers require frequent cleaning before sludge becomes too difficult to remove and reduces effectiveness. In the case of lamella clarifiers, sludge buildup on the inclined plates results in uneven flow distribution which could harm efficiency (US EPA, 2003). For this reason, maintenance requirements for lamella clarifiers are higher, but they can be reduced through the implementation of removable plates (Wilson, 2005). Clarifiers also only remove solids, so pH will not be affected, leading to the need for further pH adjustment (NMED Surface Water Quality Bureau, 2015).

=====Clarifier Design Calculations and Typical Design Values=====

======Detention Time======

Detention time (DT) is the time is takes for a unit of water to travel from the inlet of the clarifier unit to the outlet. During typical operations, the design value for this is 2 to 3 hours.

<math>\begin{align}
DT &= \frac{Tank\ Volume}{Influent\ Rate}
\end{align}</math>

======Surface Overflow Rate======

Surface overflow rate (SOR) measures the flow into the clarifier per square foot of surface area. Typical design values are 400 to 800 gal/day/sq. ft.

<math>\begin{align}
SOR &= \frac{Volumetric\ Flow\ Rate}{Surface\ Area}
\end{align}</math>

======Weir Overflow Rate======

Weir overflow rate (WOR) describes the flow in gallons per day per linear foot of weir. Typical values are 10,000 gal/day/ft.

<math>\begin{align}
WOR &= \frac{Volumetric\ Flow\ Rate}{Weir\ Length}
\end{align}</math>

======Solids Loading Rate======

Solids loading rate (SLR) describes the mass of solids in the clarifier influent per square foot of surface area. This value should not exceed 30 lbs/day/sq. ft.

<math>\begin{align}
SLR &= \frac{Solids\ Mass\ Flow\ Rate}{Surface\ Area}
\end{align}</math>

===Flotation===
Flotation is a process designed for specific solid-solid mixtures. It works by generating gas bubbles in a liquid that attach to selected solid particle. Afterwards, the particles rise to the liquid surface where they are removed by an overflow weir or mechanical scraper. The separation depends on the surface properties of the particles and its preference to attach to the gas bubbles. To meet the necessary requirements of the flotation process, a number of additives can be used to control things like the pH of the liquid-solid mixture, the activity of the solid surface, and the froth that can assist in separation. The bubbles can be produced by gaseous dispersion, dissolution, or electrolysis of the liquid (Peters & Timmerhaus, 2003).

===Centrifugation===
This process is similar to external field separation in that an external force field is applied to separate a mixture. When gravity separation is too slow due to particle densities, particle size, settling velocity, or the formation of an emulsion, centrifugation is commonly used. Centrifugal force increases the total force acting on the particle and results in faster separation times. This process is generally used to separate solids from liquids, however it can also be used to separate two liquids with very different densities (Peters & Timmerhaus, 2003).

===Drying===
Drying is performed to remove liquid from a liquid-solid mixture and produce a dry solid. Water is most often the liquid removed, but organic liquids are removed from solids on occasion as well. The heat required to vaporize the liquid is usually obtained by a series of gas-solid contacting devices. Feed condition and temperature sensitivity of the solid dictate the type of contacting device that is used. There are two groups of dryers that differ by the dependence of either mechanical means or fluid motion for gas solid contact. Another feature of dryers is to use either direct (hot gas) or indirect (conductive surface) heating (Peters & Timmerhaus, 2003).

===Evaporation===
Evaporators separate solvents from a solution by evaporation. The difference between evaporation and distillation is that evaporation requires the solute be nonvolatile. Because of this, a high separation can be achieved with one stage. Evaporators are essentially reboilers, so evaporation is a very energy-intensive process with a high thermal economy (Peters & Timmerhaus, 2003).

===Filtration===
Filtration is a process that separates a mixture of solid in a liquid or gas by passing the mixture through a porous medium in which the particles do not pass. Filtration is done by either cake filtration (particles found on the surface of the filter) or depth filtration (particles found within the filter). Cake filtration is generally performed with a cloth as the filtration medium (Peters & Timmerhaus, 2003).

==Conclusion==
Separation is a key part of most chemical processes, and there is a great variety of techniques to perform separation of compounds based on size, volatility, charge, and many other features. A common technique with which the process engineer should be familiar is distillation, but he or she should also be aware of the other available options. Some techniques may be less expensive, less energy-intensive, or more effective than distillation, depending on the specific separation problem. Therefore, the separation strategy should be carefully considered.

==References==
Air Control Technology Fact Sheet. Enviornmental Technology Fact Sheet. http://www3.epa.gov/ttncatc1/dir1/fcyclon.pdf

Belter PA, Cussler EL, Hu WS. Bioseparations: Downstream Processing for BIotechnology. New York: John Wiley; 1998.

Biegler LT, Grossmann IE, Westerberg AW. Systematic Methods of Chemical Process Design. Upper Saddle River: Prentice Hall; 1997.

Danckwerts P (1965) The Absorption of Gases in Liquids. Pure and Applied Chemistry UK 10:625-642.

Development Document for the Final Effluent Limitations Guidelines and Standards for the Metal Products and Machinery Point Source Category (Report). US Environmental Protection Agency. 2003.

Erwin, D. Industrial Chemical Process Design. New York: McGraw Hill, Professional Engineering; 2002.

Hage, David S. 1999Affinity Chromatography: A Review of Clinical Applications. Clinical Chemistry 45(5): 593–615.

Harrison RG, Todd P, Rudge SR, Petrides, DP. Bioseparations Science and Engineering. New York: Oxford University Press; 2003.

Hage, David S. 1999Affinity Chromatography: A Review of Clinical Applications. Clinical Chemistry 45(5): 593–615.

HisPur Cobalt Resin - Thermo Fisher Scientific N.d. https://www.thermofisher.com/order/catalog/product/89964, accessed February 21, 2016.

Lamella Plate Clarifier. Hydro International Web site. Available at: http://www.hydro-int.com/uk/products/lamella-plate-clarifier?s=0&r=uk. Accessed February 2, 2016.

Lean Oil Absorption. PetroGas Systems Web site. Available at: http://petrogassystems.com/technology/natural-gas-processing-and-dew-point-control/lean-oil-absorption. Accessed February 19, 2014.

Lichty, Jordan J., Joshua L. Malecki, Heather D. Agnew, Daniel J. Michelson-Horowitz, and Song Tan 2005Comparison of Affinity Tags for Protein Purification. Protein Expression and Purification 41(1): 98–105.

Merichem Gas Technologies. ®LO-CAT PROCESS available at http://www.merichem.com/images/casestudies/Desulfurization.pdf Accessed 6 Feb. 2015.

Miller L.N. & Zawacki T.S. , US 4080424, "Process for acid gas removal from gaseous mixtures", issued 21 Mar 1978, assigned to Institute of Gas Technology

NMED Surface Water Quality Bureau, New Mexico Water Systems Operator Certification Study Manual, New Mexico Environment Department, 2015.

Peters MS, Timmerhaus KD. Plant Design and Economics for Chemical Engineers. 5th ed. New York: McGraw Hill; 2003.

Prihasto, N; Lui, Q; Kim, S. Pre-treatment strategies for seawater desalination by reverse osmosis system. 2009; 249(1): 308-316. doi:10.1016/j.desal.2008.09.010

Schmidt Eberhard (2012) Waste Gases, Separation and Purification. Ullman’s Encyclopedia of Industrial Chemistry Germany 2:174-181.

Seider, W.D., Seader, J.D., and Lewin, D.R. (2004). ''Process Design Principles: Synthesis, Analysis, and Evaluation.'' New York: Wiley.

Stripping Column. Alfa Laval Web site. Available at: http://www.alfalaval.com/solution-finder/products/soft-column/Documents/Stripping%20Column.pdf. Accessed February 19, 2014.

Towler G, Sinnott R. Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013.

Turton, R.T., Bailie, R.C., Whiting, W.B., and Shaewitz, J.A. (2003). ''Analysis, Synthesis, and Design of Chemical Processes'' Upper Saddle River: Prentice-Hall.

Wankat, P.C. (2012). ''Separation Process Engineering.'' Upper Saddle River: Prentice-Hall.

Wilson, T.E., Clarifier Design, 2nd Ed., McGraw-Hill: New York, 2005.

Separation processes

2016-02-22T01:30:11Z

Bts511: /* References */

Authors: Nick Pinkerton, [2014] Karen Schmidt, [2014] James Xamplas, [2014] Emm Fulk, [2015] and Erik Zuehlke, [2015] John Dombrowski [2016] , Brett Sleyster [2016] , Robert Cignoni [2016] , and Osman Jamil [2016] 

Stewards: David Chen, Jian Gong, and Fengqi You

Date Presented: February 9, 2014 /Date Revised: February 1, 2014

 

==Introduction==
Essentially all chemical processes require the presence of a separation stage. Most chemical plants comprise of a reactor surrounded by many separators. Separators have a countless number of jobs inside of a chemical plant. A separator can process raw materials prior to the reaction, remove incondensable gases, remove undesired side products, purify a product stream, recycle materials back into the process, and many other jobs that are essential to the process.

Chemical engineers must understand the science of separation and the variety of ways that separation can take place. There are many ways to perform a separation some of these including: distillation, absorption, stripping, and extraction. The science of separation revolves around the presence of two phases that are in contact and equilibrium (Wankat, 2012).

[[File:Sepmeth.JPG|frame|Figure 1. Separation methods by property]]

==Theory==
===Vapor-Liquid Equilibrium===
Separation processes are based on the theory of vapor-liquid equilibrium. This theory states that streams leaving a stage in a separation process are in equilibrium with one another. The idea of equilibrium revolves around the idea that when there is vapor and liquid in contact with one another they are in constantly vaporizing and condensing. Different components in the mixture will condense and vaporize at different rates. There are three types of equilibrium conditions that can be subdivided into thermal, mechanical and chemical potential categories. These separate equilibrium states are given as:

<math>T_{liquid} = T_{vapor}</math>

<math>p_{liquid} = p_{vapor}</math>

<math>chemical potential_{liquid} = chemical potential_{vapor}</math>

==Distillation==
===Flash Distillation===
Flash Distillation is one of the simpler separation processes to be employed in a chemical plant. The main premise of flash distillation is that a portion of a liquid feed stream vaporizes in a flash chamber or a vapor feed condenses. Vapor-liquid equilibrium will cause the vapor phase and the liquid phase to have different compositions. The more volatile component of the mixture will compose of a larger portion of the vapor. This simple separation is easy to manufacture but does not result in large degrees of separation.

Flash distillation requires a feed stream that is pressurized and heated and then passed through a valve into a flash drum. The large pressure drop across the valve will result in a partial vaporization of the fluid. Vapor will be removed overhead from the flash drum while the remaining liquid will collect at the bottom of the drum and be removed. Most flash drums will contain an entrainment eliminator which is a screen that prevents liquid from being carried into the vapor effluent. Figure 2 shows a simple overview of the flash distillation process. As shown, there is a heater that flows into a let-down valve where the two-phase flow begins. Variables y and x are the mole fractions of the more volatile component in the vapor and liquid effluents, respectively.

[[File:Flash.gif|center|frame|Figure 2. Flash Distillation Flow Diagram]]

===Column Distillation===
Distillation columns are the most widely used separation technique used in the chemical industry, accounting for approximately 90% of all separations (Wankat, 2012). Distillations in columns consist of multiple trays that each act at their own equilibrium conditions. Large columns are able to perform complete separations of binary mixtures as well as more complex multi-component mixtures.

[[File:column.jpg|250px|center|]]
===Stages===
Columns are separated into stages by the presence of trays. These trays allow for vapor-liquid contact and equilibrium to occur. Typically, the more stages in a column, the larger separation that can be achieved. There are many different types of trays that can be used in a column.
====Sieve Trays====
The simplest and least expensive tray type is the sieve tray which is a sheet of metal with holes punched into it to allow vapor flow. Sieve trays can have different hole patterns and sizes that will affect the tray efficiency and flow rates.

[[File:sieve.jpg|200px|center|]]

====Sieve Tray Design Procedure====

The design of these plates is done through a trial-and-error process. Most commercial process simulations (such as HYSYS) have default tray designs, and automatically specify dimensions. However, these dimensions selected or calculated by the simulations may not give the best performance for your system, so it is valuable to understand how to design the sieve trays and how specific parameters may affect performance. Hand calculations using the following methods can be used to guide the simulation programs to better design. This section will use sample data to work through an example of the process. The following is a general list of steps for designing a sieve plate:

=====1. Calculate the maximum and minimum vapor and liquid flow rates for the turndown ratio required.=====
This data can be collected from a McCabe-Thiele diagram and/or from process simulation data.

Data from McCabe Thiele diagram, for example: 
Number of stages = 10 
Slope of top operating line = 0.185 
Slope of bottom operating line = 1.43 
Top composition = 98.8 mol% acetone 
Bottom composition = 4 mol% acetone (a) 
Minimum reflux ratio = 0.31 

Collect data rom mass balances: 
MW of feed = (0.6 kmol a/kmol)*(58 kg/kmol a) + (0.4 kmol acetic acid/kmol)*(60 kg/kmol aa) = 59 kg/kmol 
Feed rate = 100 kmol/h 
Using material balances... D = 57.3 kmol/h ; B = 42.7 kmol/h 
Vapor rate, V = D(1+R) = 57.3 kmol/h (1+.31) = 75.1 kmol/h

=====2. Collect or estimate the system physical properties.=====
Here it is important to know information about both the top and bottom of the column. Useful information includes temperature, pressure, column pressure drop (a common assumption is 100 mmH2O per plate), densities, molecular weights, surface tensions, and number of stages (which can be estimated from the McCabe-Thiele diagram).

=====3. Select a Trial Plate Spacing=====
The plate spacing will depend on the column diameter and operating conditions. Plate spacings from 0.15 m to 1.0 m are typically used. The smaller the diameter, the smaller the spacing. Small columns will use close spacing. Columns with diameters above 1.0 m, plate spacings of 0.3 m to 0.6 m are normally used. A good initial estimate is 0.5 m.
[[File:trayspacing.jpg|400px|center|]]

=====4. Estimate the column diameter, based on flooding considerations.=====
Vapor and liquid flow rates will vary along the column, so plate design needs to be considered both above and below the feed. Using plate spacing and FLV (which is the square root of the ratio of the liquid to vapor flow rates), you can obtain the value of K from the plot.

[[File:floodingplot.jpg|400px|center|frame|Figure. Plate Spacing]]
There is a range of vapor and liquid flow rates in which the column needs to be operated. Too low or too high of rates can result in various inefficiencies in the column operation, as shown in the figure below. For example, if the vapor rate is too high, flooding will occur. However, it is not safe to operate on the flooding line. Instead, columns are typically designed for 80% of flooding at the maximum flow rate.
[[File:vap_rate_vs_liq_rate.jpg|400px|center|frame|Figure. Tray behavior]]

=====5. Decide the liquid flow arrangement.=====
Common flow arrangements are single pass (cross flow), double pass, and reverse flow. Using conditions at the bottom of the column, calculate the max volumetric flow rate. Use this flow rate and the column diameter to determine the preferred flow arrangement from the chart below.
[[File:Liquidflow.jpg|200px|center|]]

=====6. Make a trial plate layout: downcomer area, active area, hole area, hole size, weir height.=====
Standard sizes for trays -- and good assumptions for the first iteration -- are: weir height, hw = 50mm ; hole diameter, Dh = 5mm ; plate thickness, tpl = 5mm. From the graph below, the ratio of downcomer area (Ad) to column cross-sectional area (Ac) can be determined from the ratio of weir length (lw) to column diameter (Dc) and vice versa.
[[File:platelayout.jpg|200px|center|frame|Figure. Plate Dimensions]]

=====7. Check the weeping rate=====
Compare the actual vapor velocity to the minimum vapor velocity -- if velocity is too low fluid will "weep" through the tray holes. If the weeping rate is unsatisfactory, return to step 6 and choose different values for the plate layout dimensions. From the chart in step 4, it can be seen that there is a minimum vapor flow rate below which the liquid "weeps" from the tray above.

For the remaining steps in this design process, it is recommended to check your assumptions after each step and revise them as necessary in order to maintain operation in the "sweet spot" of the vapor rate vs. liquid rate plot. Additional iterations may be required as you move through the procedure.

Calculate the maximum liquid flow rate. Calculate the minimum liquid flow rate at 70% turndown (recommended). Calculate the height over the weir as
<math>h_o=750[\frac{L_w}{p_Ll_w}]^\frac{2}{3}</math>

=====8. Check the plate pressure drop=====
<dfn>If the pressure drop calculated here is too high, return to step 6.</dfn>

Proceed to step 9 if the pressure drop assumption is valid.
=====9. Check the downcomer backup. =====
<dfn>If the downcomer backup is too high, return to step 6 or 3.</dfn>

The plate spacing affects the amount of fluid in the downcomer. Calculate the level in the downcomer and the residence time of the fluid to see if the values are valid. Note that residence times greater than 3 seconds are acceptable.

Proceed to step 10 if residence time is acceptable.
=====10. Decide plate layout details.=====
Determine calming zones, the unperforated areas at the inlet and outlet sides of the plate. The width of each zone is usually made the same. Recommended values are: below 1.5 m diameter, 75 mm; above, 100 mm. The unperforated area can be calculated from plate geometry. Also check the hole pitch, or the distance between hole centers. It should not be less than 2.0 hole diameters. A normal range is between 2.5 and 4.0 hole diameters. The shape must also be specified. Square and equilateral triangle holes are used.

=====11. Recalculate the percentage flooding based on the chosen column diameter.=====
An assumption of 80% flooding was chosen so that operation would occur in the "sweet spot." This assumption must be checked by calculating the flooding percentage for a given column diameter.
uv = (max volumetric flow rate)/(net area)

<math>%flooding = \frac{u_v}{u_f}</math>

If the hole pitch is unsatisfactory, return to step 6.
=====12. Check entrainment=====
''If too high, return to step 4'' Use the graph below to determine entrainment from FLV.
[[File:entrainment.jpg|400px|center|]]

The value for fractional entrainment can be used to re-estimate the column efficiency, and reevaluate the number of trays needed. Can return to step 1 for more accurate estimates.

=====13. Optimize design.=====
After returning to step 1 to reevaluate the number of trays, it is valuable to repeat steps 2 through 12 to find the smallest diameter and plate spacing acceptable at the lowest cost.

=====14. Finalize the design.=====
Optional: draw up the plate specification and sketch the layout of the plate.

====Bubble-Cap Trays====
Bubble-cap trays consist of a weir around each hole in the tray which is covered with a cap that has holes or slots to allow vapor passage. Entrainment is about three times larger than a sieve tray. Bubble-cap trays require larger tray spacing than sieve tray design. Bubble-cap trays have been known to have problems with coking, polymer formation, or high fouling mixtures. Recently, very few new bubble-cap columns are being built due to the expense and marginal benefits. However, engineers will likely encounter bubble-cap columns still currently in operation.

====Flow Patterns====
Cross flow columns are the most common pattern for distillation columns. For liquid flows between 50 and 500 Gal/min, a cross flow column is appropriate. When liquid flow is increased above 500 Gal/min, an engineer should consider designing a double pass or multi-pass column. This will reduce the liquid gradient on the tray and reduce the downcomer loading (Wankat, 2012).

===Column Sizing===
Column height will be dependent on the amount of trays required and the spacing between the trays. Normally, tray spacing of 0.15 m to 1 m is used. For columns, above 1 meter in diameter, 0.5 m can be used as an initial estimate.

Column diameter is influenced by the vapor flow rate in the column. The trays can not have excess liquid entrainment or high pressure drops; therefore, vapor velocity in the column must be maintained at a reasonable level.

An equation based on the Souders and Brown equation can be used as an estimate for the max allowable superficial vapor velocity,

<math>\hat u_v = (-0.171l_t^2 + 0.27l_t - 0.047){\frac{\rho_L - \rho_v}{\rho_v}}^{1/2}</math>

where <math>l_t</math> is the plate spacing in meters, <math>\rho_L</math> is the density of the liquid stream, and <math>\rho_V</math> is the density of the vapor stream.

Column diameter, <math>D_c</math>, can then be estimated using the relation,

<math>D_c = \sqrt{\frac{4\hat{V_w}}{\pi\rho_v\hat{u_v}}}</math>

where <math>\hat{V_w}</math> is the maximum vapor rate in kg/s (Towler et al., 2013).

===Distillation Applications===

Distillation is a process that can be implemented in various scales. There is both laboratory scaled distillation as well as very large industrial distillation. Other applications for distillation include food/alcohol processing and herb distillation for the perfume and medical industries. Typically laboratory scaled distillation occurs in batches whereas industrial distillation (e.g. fractional distillation of crude oil) occurs continuous with a constant distillate and bottom effluent streams.

Some applications of distillation are concerned the top stream only, some the bottom stream only and others both streams can be used for future products. In alcohol distillation for example, the water that is separated from the ethanol/water binary solution is discarded as waste water. In fractional distillation of crude oils, the heavy hydrocarbons at the bottom of the column are collected and sold along with the light hydrocarbons that appear in higher side draws (Wankat, 2012).

===Example Case: Ideal Distillation===

Assume an equimolar mixture flowing at 10 mol/s of 20 mol% n-pentane, 30 mol% n-hexane, and 50 mol% n-heptane. Separate the mixture into 3 products: 99% pure n-pentane, 99% pure n-hexane, 99% n-heptane. Assume the feed and products are all liquids at the bubble points. There are two process alternatives to consider in this example. The direct sequence removes the most volatile species, pentane, in the first column, and then separates hexane and heptane in the second column. The indirect sequence separates the heaviest product, heptane, and then separates pentane from hexane in the second column. This example will consider the direct sequence. Next, we must decide if these species exhibit fairly ideal behavior during distillation. Since the n-alkanes have very similar properties, it is safe to assume they will display close to ideal behavior. The next step is to look up the boiling points of the 3 species. In this case, the normal boiling points of pentane, hexane, and heptane are 309 K, 342 K, and 372 K, respectively. Also, it is a good idea to look up relative volatilites, to further verify near-ideality of the mixture, but also to obtain the information necessary for the Underwood method, which we will employ to obtain a solution. The next step is to write out material balances based on molar flows and the design specifications. They go as follows:

<math>\mu_I(nC5) + \mu_{II}(nC5) = 2 mol/s</math>

<math>\mu_I(nC6) + \mu_{II}(nC6) + \mu_{III}(nC6) = 3 mol/s</math>

<math>\mu_{II}(nC7) + \mu_{III}(nC7) = 5 mol/s</math>

<math>\mu_I(nC5) = 99\mu_I(nC6)</math>

<math>\mu_{II}(nC5) = (5/990)\mu_{II}(nC6)</math>

<math>\mu_{II}(nC7) = (5/990)\mu_{II}(nC6)</math>

<math>\mu_{III}(nC7) = 99\mu_{III}(nC7)</math>

where <math>\mu</math> represents the molar flow, and the subscript represents the product stream.

Solving this system of equations:

<math>\mu_I(nC5) = 1.985\ mol/s</math>

<math>\mu_{II}(nC5) = 0.015\ mol/s</math>

<math>\mu_I(nC6) = 0.020\ mol/s</math>

<math>\mu_{II}(nC6) = 2.930\ mol/s</math>

<math>\mu_{III}(nC6) = 0.050\ mol/s</math>

<math>\mu_{II}(nC7) = 0.015\ mol/s</math>

<math>\mu_{III}(nC7) = 4.985\ mol/s</math>

At this point we have enough information to use Underwood's method to estimate the minimum vapor flows in the column. The following three equations are used in Underwood's method:

<math>\sum_i \frac{\alpha_{ik}}{\alpha_{ik}-\phi}f_i = (1-q)F</math>

<math>(R_{min}+1)D = \sum_i \frac{\alpha_{ik}}{\alpha_{ik}-\phi}d_i = V_{min}</math>

<math>\bar R_{min}B = -\sum_i \frac{\alpha_{ik}}{\alpha_{ik}-\phi}b_i = \bar V_{min}</math>

where <math>\alpha_{ik}</math> is the relative volatility of species <math>i</math> to species <math>k</math>, <math>f_i</math> the molar flow of species <math>i</math> in the feed, <math>q</math> the fraction of the feed that joins the liquid stream at the feed tray, <math>F</math> the total molar flow of the feed, <math>D</math> the molar flow of the distillate, <math>R_{min}</math> the minimum reflux ratio <math>(=L_{min}/D)</math>, <math>d_i</math> the molar flow of species <math>i</math> in the distillate, <math>V_{min}</math> the minimum vapor flow possible in the top section of the column to accomplish the desired separation, <math>\bar R_{min}</math> the minimum reboil ratio <math>(=\bar V_{min}/B)</math>, <math>b_i</math> the molar flow of species <math>i</math> in the bottoms product, and <math>\bar V_{min}</math> the minimum vapor flow in the bottom section of the column. The final variable, <math>\phi</math>, will be solved for using the first Underwood equation, and it's value will be decided based on the relative volatilities of the key components in the column.

So, after solving the first Underwood equation, we get two values for <math>\phi</math>, 3.806 and 1.462. Because 3.806 is between the relative volatilities of the key components, we will substitute that value for <math>\phi</math> into the second Underwood equation. Doing so for both columns gives <math>V_{min} = 6.4\ mol/s</math> for the first column and <math>V_{min} = 8.9\ mol/s</math> for the second column, for a total minimum vapor flow of 15.3 mol/s. The process would then be repeated for the indirect sequence, and the decision for which process to use would be justified by the process with the overall minimum vapor flow (Biegler et al., 1997).

==Absorption==
===Description of Absorption===
Another separation process used in industry is absorption, which is used to remove a solute from a gas stream. It accomplishes this by contacting the gas mixture with a liquid solvent that readily absorbs the undesirable components from the gas stream, purifying the gas stream. This separation process is determined by the inputs of the liquid flow rate, temperature, and pressure.

The absorption factor, which can be determined mathematically, determines how readily a component will absorb in the liquid phase. The absorption factor of component i is

<math>A_i=L/K_iV</math>

where <math>L</math> is the liquid flow rate entering the column, <math>V</math> is the vapor flow rate entering the column, and <math>K_i</math> is the vapor/liquid equilibrium ratio for component i (Peters & Timmerhaus, 2003). Higher absorption factors result in higher absorptivity into the liquid and a decrease in the number of trays required for separation, however a diminishing return occurs after the absorption factor is greater than 2.0. An absorption factor of 1.4 is most commonly used.

In general absorption can be seperated into two overarching categories, physical and chemical absorption. In physical absorption, the unwanted solute in the gas is absorbed into the liquid phase because solubility of the component is higher in the liquid phase than the gas phase. In chemical absorption the solute is removed from the gas via a reaction with the solvent, this reacted product is then transported into the liquid phase (Danckwerts 1965). There are two types of chemical absorption reversible and irreversible. Generally reversible chemical absorption is preferred as the solvent can be put through a stripper and regenerated so it can be recycled back to the absorption process (Wankat, 2012).

===Absorption Apparatus===

There are five major apparatus used for absorption in industrial application. These five pieces of equipment are spray absorbers (or towers), ejector (venturi) scrubbers, packed columns, trayed columns, and film absorbers (Schmidt, 2012).

==== Spray Tower vs Ejector Scrubber ====

In both '''spray tower''' and the '''ejector scrubber''' nozzles are employed to produce small solvent droplets. These small droplets increase the surface area of the liquid to gas contact allowing for the maximum amount of mass transfer to occur between the gas mixture and the liquid. The major difference between the two nozzle equipment designs is the configuration and type of nozzles. In the ejector scrubber shown in Figure 3 there is a single nozzle that is generally a higher pressure spray nozzle that produces finer solvent drops allowing for an even greater amount of mass transfer enabling better physical absorption (Schmidt, 2012).
[[File:Ejectorventuri.jpg|thumb|200px|center|Figure 3. Ejector Scrubber (US EPA, 2006)]]
'''Spray towers''' on the other hand generally have many nozzle at different heights where the liquid solvent will be sprayed out of to contact the gas running through the tower. This design is used in order to ensure the gas contacts the liquid as throughout the tower. These nozzles are lower pressure than a ejector scrubbers nozzle and thus physical mixing is worse in this configuration. Since physical mixing is generally worse in this configuration it is usually used in conjunction with a chemical absorption process. The other major difference between the ejector scrubber and the spray tower is that gas and liquid flow is cocurrent in the former while it is countercurrent in a spray tower. A spray tower absorber is shown below in Figure 4 (Schmidt, 2012).
[[File:SparyTowerAbsorber.jpg|thumb|200px|center|Figure 4. Spray Tower Absorber (US EPA, 2006)]]

==== Tower Type Absorption Apparatus ====
'''Packed column absorbers''' and '''tray column absorbers''' have very high efficiencies for the removal of an unwanted solute in the gas stream. The major disadvantage a trayed column has when compared to a packed column is the pressure drop. The pressure drop in a packed column is generally very low, whereas in between each tray of a trayed column pressure drop can be quite large. However the advantages inherent to trayed columns become clear when one needs the solvent to have a high concentration of the component to be removed from the gas stream. This is most important in the case where there is a very low concentration of the component in the gas stream and the specification states the solvent must contain a high concentration of that component. In this case the flow rate of the solvent may not be high enough for a packed column, however in a trayed column the solvent flow rate can be near zero for operation (Schmidt, 2012). Packed and trayed column internals are very similar to the setups found in the respective distillation columns.

For a '''trayed column''' the plate efficiency can be calculated using O'Connell's Correlation which invovles the Henry's Law constant, total system pressure, and solvent viscosity at the operating temperature (Towler & Sinnott, 2013).

<math>x=0.062*\frac{\rho_s*P}{\mu_s*H*M_s}</math>

where
<math>x</math> is the tray efficiency,
<math>\rho_s</math> is the density of the solvent in <math>kg/m^3</math>,
<math>P</math> is the total pressure of the system in <math>N/m^2</math>,
<math>\mu_s</math> is the solvent's viscosity in <math>mNs/m^2</math>,
<math>H</math> is the Henry Law constant in <math>1/(Nm^2*(mol fraction))</math>,
and <math>M_s</math> is the molecular weight of the solvent.

A packed towers height can be determined using the equations below when concentration of solute is below 10% so that the assumption that the flow of gas and liquid will be essentially constant throughout the column holds (Towler & Sinnott, 2013). The height of packing <math>Z</math> is given by the following equation:

<math>Z=\frac{L_m}{K_G*a*P}*\int\limits_{y_2}^{y_1} \frac{dy}{y-y_e}\,</math>

where <math>P</math> is the total pressure, <math>a</math> is the interfacial surface area per unit volume, <math>y_1</math> and <math>y_2</math> are the mol fractions of the solute in the gas stream at the bottom and top of the column respectively, <math>G_m</math> is the molar gas flow rate per unit cross-sectional area, and <math>y_e</math> is the mole fraction of solute in the gas that would be in equilibrium with the liquid concentration.

The first half of the equation before the integral can be called the height of an overall gas-phase transfer unit <math>H_G</math> and the second part of the equation is the number of overall gas-phase transfer units or <math>N_G</math>. Using these definitions the above equation can be simplified to

<math>Z=H_G*N_G</math>

These equations assist in sizing an absorption column (Towler & Sinnott, 2013).

==== Film Absorber ====
The final absorber the film absorber is generally used in the case where the heat of absorption must be removed. The film absorber operates by sending the gas and solvent through a heat exchanger where the solvent creates a thin film on the walls of the tubes and the gas flows through the interior allowing for solute transfer. The good heat transfer present in a film absorber makes it preferable for situations where low temperatures are required for a high recovery of the solute (Schmidt 2012).

===Industrial Absorption Processes===
An industrial example is lean oil absorption, which is used to separate nitrogen and other impurities from natural gas. A lean oil is contacted with low quality natural gas, and the methane is selectively absorbed by the lean oil, leaving the impurities behind. The methane is subsequently regenerated from the rich oil as high quality natural gas (Petrogas Systems, 2014).

Other common industrial practices of absorption come from sour gas treatment. Amine gas treating is used to remove hydrogen sulfide or carbon dioxide from gas streams via a reversible chemical absorption. In amine gas treating the sour gas is fed to the bottom an absorber where amine solution is fed to the top along with any necessary make up water. The sour gas components are absorbed into the amine via a chemical absorption method. Sweet gas leaves the top of the absorber whereas the amine out of the bottom, now rich with acidic components is sent to a regenerator where the acid gas components are stripped and the acid gas is generally sent to a flare whereas the amine now lean again is recycled back into the first absorber (Miller & Zawacki, 1978). Figure 5 below shows the typical setup of an amine plant. Another type of sour gas treatment that uses absorption is Merichems LO-CAT process which uses a chelated iron to remove hydrogen sulfide from feed gas in the absorption column (Merichem 2015).
[[File:AmineTreating.png|thumb|400px|center|Figure 5. Amine Gas Treating Plant Schematic]]

==Stripping==
This process separates solutes from solvents (often after absorption, to purify the solvent so that it can be recycled to an absorber). Stripping will depend on the vapor and liquid flow rates, as well as the temperature and pressure of the column. There is a temperature drop down the column, so columns generally have either an increased operating temperature or decreased operating pressure.

The stripping factor of component i is

<math>S_i=K_iV/L</math>

where <math>K_i</math> is the vapor/liquid equilibrium ratio, <math>V</math> is the vapor flow rate entering the column, and <math>L</math> is the liquid flow rate entering the column, will determine how much of solute i will be stripped from the liquid into the vapor phase (Peters & Timmerhaus, 2003). The usual range for the stripping factor is between 1.2 and 2.0, with a stripping factor of 1.4 being most economic.

An example of stripping in industry is the deodorization of food items such as oils. The oil is heated and allowed to trickle down the column while steam flows up from the bottom of the column. At the vapor-liquid interface, volatile components of the oil transfer to the steam and are carried off the top of the column, leaving a purified oil product (Alfa Laval, 2014).

==Bioseparations==
===Importance===
As our ability to manipulate and engineer biological systems improves, biological products are becoming an increasingly important source of therapeutics and fuels. The production of fuels from biomass via either the enzymatic breakdown of a feedstock or the secretion of usable lipids from algae is a promising new energy source. Additionally, enzymes, antibodies and other therapeutic proteins have been applied to the treatment of a wide range of diseases. Although each process requires its own set of separations, all follow the same basic format: separation of biomass, product isolation, and product purification (Belter et al., 1998). This section will provide examples of unit operations in each step. Ultimately, the choice of separation process and unit operations will depend on the specific process and product. The descriptions below are examples of the most common bioseparation operations within the general platform (Harrison et al., 2003).

Bioprocesses begin with fermentations or growth operations. In biofuel production processes, this may involve growing algae or breaking down corn or cellulosic biomass. For the production of therapeutics, mammalian or bacterial cells may be grown in a fermentor and the product secreted into the supernatant or harvested from the cells.

===Biomass Separations===
After fermentation and product production, the solid biomass must first be separated from the desired product. If the product is secreted from the cells, this can be done immediately after fermentation ends. If the product is not secreted, the cells must first be lysed.
Cell lysis is the process of lysing, or breaking, the cell in open. Mechanical lysis is the simplest, and involves physically breaking the cell either by mashing (think mortar and pestle) or blending the cells into a homogenous solution in a homogenizer. Chemical lysis is another method, achieved by introducing an osmotic shock or chemically degrading the cell membrane. Additional separation can be achieved by flocculation, which is the process of aggregating biomaterial by charge neutralization or bridging. These larger complexes are easier to separate from smaller molecules (Harrison et al., 2003).

The next step is removing the unwanted biomass from the product in solution. Separation by centrifugation or sedimentation are the most common, although filtration is sometimes also used for processes where a biomass cake is desired. Both methods utilize density differences to separate the product from the solid biomass (Towler and Sinnott, 2013).

====Sedimentation====

Sedimentation relies purely on the force of gravity, while centrifugation speeds the settling process by subjecting the cells to a centrifugal force. Sedimentation in a settling tank is the simplest method of solid-liquid bioseparation. In this process, biomass in a tank is simply allowed to settle to the bottom over time. While this process is inexpensive, requires little energy and can separate out large volumes of biomass, it generally requires long time periods and is only mostly in very large-scale processes where active centrifugation is difficult (Belter et al., 1998).

====Centrifugation====
Centrifuges are widely utilized across many processes, and thus a wide variety of scales and designs have been developed. Disk-stack centrifuges, in which the solid phase is deposited onto “shelves” in the center of the spinner and liquid phase is pushed to the outside, are some of the most commonly used centrifuges in industry. They are especially suited to biomass separation processes because they can be built on a large scale and are ideal for separating fine solids from liquids. [[File: Disk_stack_centrifuge_towler.png|frame|center|Fig. 6: Diagram of a disk-stack centrifuge (Tolwer et al, 1997).]] Tubular bowl centrifuges are also common and can reach separation efficiencies of up to 90%. Heavier products accumulate along the sides of the bowl, while the light phase flows out the top. They separate products by can be used both to separate solids from liquids and immiscible liquids, such as and oil product and an aqueous broth (Tolwer and Sinnott, 2013). [[File: tubular bowl centrifuge towler.png|frame|center|Fig. 7: Diagram of a tubular bowl centrifuge centrifuge (Tolwer and Sinnott, 2013).]]

Centrifugation scale-up is made easier by sigma analysis, which allows for the estimation of appropriate feed rates for different size centrifuges. The sigma factor is dependent on the inner and outer radius of the centrifuge, the angular velocity, and the sedimentation velocity of the solid particles being separated. It can be thought of as the characteristic cross-sectional area with units of [length]2. The sedimentation velocity can be calculated by

<math>v_g={\frac{2a^2(\rho-\rho_0)}{9\mu}}g</math>

where <math>v_g</math> is the sedimentation velocity, <math>a</math> is the cell or biomass particle diameter, <math>\rho</math> is the particle density, <math>\rho_0</math> is the fluid density, and <math>\mu</math> is the fluid viscosity. The volumetric flow <math>Q</math> can be estimated by

<math>Q=(v_g)(\Sigma)</math>.

The equality

<math>{\frac{\Sigma_1}{\Sigma_2}}={\frac{Q_1}{Q_2}}</math>

can be an easy way to estimate equivalent flow rates between a small-scale centrifuge 1 and larger centrifuge 2 (Harrison et al., 2003).

====Example: Centrifugation Scale-up====

You are trying to separate a cell of radius 0.4 <math>\mu</math>m with a density of 1.05 g/cm3 from broth of mostly water (density of 1 g/cm3 and viscosity of 0.01 g/cm s). The sigma factor of the centrifuge you are using is 1 x 106 cm2. A] What volumetric flow rate should you use? B] If you want to scale up the process to a centrifuge with <math>\Sigma</math> = 3 x 106 cm2, what flow rate would you use in the larger centrifuge?

Solution:
A] Using the equation for <math>v_g</math>, and being mindful of units, the sedimentation velocity equals 1.74 x 106 cm/s. The flow rate, then, equals

<math>Q=(1.74 x 10^-6)(1,000,000) = 1.74 cm^3/s = 0.104 L/min</math>.

B] Keeping in mind that for the same process, <math>v_g1 = v_g2,</math> and rearranging the sigma factor equality, the new flow rate is

<math>Q_2 = {\frac{\Sigma_2 x Q_1}{\Sigma_1}} = {\frac{(3 x 10^6)(0.104)}{1 x 10^6}} = 0.313 L/min </math>

===Product Isolation===
Liquid-liquid separation, to extract the product from the aqueous phase, is much less straightforward than liquid-solid extraction. Many methods - especially adsorption, filtration, and precipitation - are similar in principle to operations found in other, non-biological separations. The exact separations used depend on the nature of the product and the scale of the process. These processes are nearly identical to their non-biological counterparts, and their description is left to other sections.

Particular care needs to be taken with protein products because of their instability, and the selection of an appropriate solvent or adsorbent is crucial to a successful process (Harrison et al., 2003).

===Product Purification===
The final steps of protein purification and polishing remove any remaining contaminants and bring the concentration of product to an appropriate value for applications. Purification processes for food-grade and medical products can be extensive, as sterility and high purity are essential. Purification in fuel-producing processes may be less extensive, depending on the process. Chromatography and crystallization are two common steps in purification and are especially used in industrial scale protein production. Several different types of chromatography exist with the ability to carry out different types of separations.
Chromatography is similar to adsorption in that it relies on differences in affinity between solutes and a solid surface. A solution is eluted through a column containing a solid resin with various affinities for the substances in solution. In adsorption, the solutes are evenly saturated throughout the column. Chromatography differs in that solutes are deposited a resin phase before the column is flushed with an elution solvent specific that results in solutes eluted in bands.

==== Ion Exchange Chromatography ====

There are two main types of ion exchange columns—anion and cation. Anion exchange resins have a positive charge and are used to retain products with a negative charge. Cation exchange resins have a negative charge and are used to retain products with a positive charge. The pH of the elution buffer is change to force a specific solute to wash out, depending on whether the pH of the buffer is above or below the isoelectric point of the solute (Belter et al., 1998). This is especially useful for the separation of protein product (including antibodies), nucleic acids, and other charged molecules. When the solutes have sufficiently different isoelectric points, the pH of the buffer is manipulated to affect the solute charge and force the product to elute while the solute remains preferentially bound to the resin, or vice versa (Harrison et al., 2003). In general, the most strongly charged molecules will remain in the column for a longer period of time. Elution washes through the weakly bound ions before the more strongly bound ions. Different speeds of elution can be visualized as in figure 8.

[[File:chromatography.png|frame|center|Fig. 8: Illustration of product bands in an elution chromatography column (Belter et al., 1998).]]

==== Size Exclusion Chromatography ====

In gel filtration chromatography, small molecules are "trapped' by the porous resin and take longer to flow through the column. Larger products will elute first because the smaller molecules are better able to penetrate the resin. This forces them to take a much longer path through the column, which means it takes longer for them to elute. This operation is often used when there is a distinct difference in size between the desired product and other solutes.

==== Affinity Separations ====

Affinity chromatography is very similar to ion exchange chromatography in that the interactions between the material in the column and the molecules in the feed. The main difference is that affinity chromatography can rely on a great variety of types of interactions. Two very common types of affinity are exploited in affinity chromatography columns. The first is immunoaffinity. Proteins are specifically bound by antibodies which can be incorporated onto beads and used in chromatography. Antibodies are designed to bind only a single protein, so these interactions are considered to be highly specific. The protein can be eluted using a buffer that changes the pH or salinity in the column, which adversely affects binding (Hage, 1999).

Proteins can be separated from very complex and unrefined mixtures using this method because of the great specificity. Antibody-protein interactions are so specific that individual proteins can be isolated from blood samples as shown in the figure below. The main drawback of this method is simply that specific interactions between proteins and binding targets do not always exist. Antibodies exist for many proteins, but for others they must be customized or created using some sort of evolution process. This is not a trivial task and can require a significant capital investment (Hage, 1999).

[[File:Affinity_Chromatography_Example.PNG|frame|center|Fig. 9: Affinity purification separating fibrinogen from human plasma using an anti-fibrinogen antibody (Hage, 1999).]]

The other main type of affinity chromatography is based on protein specific tags and the molecules or surfaces to which they bind. One of the most common types of protein tags used is the polyhistidine tag. This tag consists of 6-8 consecutive histidine residues which can be added to the exterior of the desired protein product. The addition of this tag requires alterations to the coding sequence of the protein. The polyhistidine tag binds strongly with nickel and cobalt ions. The product with the tag can then be eluted with imidazole—a small molecule with the same structure as the functional group of the amino acid histidine. Imidazole will bind the cobalt and nickel ions more strongly than the histidine in the tag. Along with chromatography, protein tag interactions can be leveraged with the use of beads that can be deposited directly into the solution containing the protein of interest (Lichty et al., 2005).

Different proteins can be purified in this manner with varying levels of efficiency. There is a very high dependence on the size of the protein, electronic properties, and steric considerations. The figure below demonstrates varying separation efficiencies with proteins of different sizes (HisPur Cobalt Resin - Thermo Fisher Scientific, 2016).

[[File:HisTagPurification.PNG|frame|center|Fig. 10: Protein gel demonstrating the separation of proteins from a sample. The far left column represents the most efficient separating conditions using cobalt and the far right column represents a negative control with no purification performed (HisPur Cobalt Resin - Thermo Fisher Scientific, 2016).]]

Several other types of tag-bead interactions can be utilized in separations processes. Maltose Binding Protein is a small protein that can be added to a protein of interest. It binds strongly with beads coated in immobilized maltose and can be released by flushing with maltose. As MBP is a full sized protein that typically must be removed from the protein of interest in order for it to be used. In this case, the site specific TEV protease is often used cleave MBP from the protein of interest. In addition, under specific circumstances, other unique tags can be used and provide varying levels of specificity in separations. The Flag tag, 3x Flag tag, Glut tag, and Strep tag. While these are all commonly used, the polyhistidine tag is the most popular because it gives the highest level of specificity. Some of the main purification tags are compared in the table below (Lichty et al., 2005).

{| class="wikitable"
|-
! Tag
! Size (aa)
! Resin
! Eluting Agent
! Capacity (mg/ml)
! Cost/10 mg
|-
| His
| 6
| Talon
| Imidazole
| 5-14
| $18
|-
| Maltose Binding Protein
| 396
| Amylose
| Maltose
| 3
| $12
|-
| Glutathione S-Transferase
| 218
| GSH-Sepharose
| Glutathione
| 10
| $36
|-
| Strep II
| 8
| Strep-Tactin Sepharose
| Desthiobiotin
| 0.1
| $293
|-
| FLAG
| 8
| Anti-FLAG M2 MAb Agarose
| Flag peptide
| 0.6
| $1045
|}

Unfortunately, based on the contents of the solution, the characteristics of the protein, and the specific product being used, there are extremely large amounts of variation in yield and purity for each type of tag. Therefore, it is crucial to allocate time towards preliminary tests and scale-up experiments with the exact solution from which the protein of interest is being separated. The his tag is typically the most used because, as seen in the table above, it has one of the lowest costs and highest yields. In addition, because it is only a 6 amino acids, it often is not cleaved from the protein of interest which minimizes the number of steps required in the overall purification process (Lichty et al., 2005).

==== Crystallization ====

Crystallization, or the formation of solute crystals from a solution, is especially useful in biomolecule separations because it is possible to obtain a 99.9%+ product purity. In crystallization, a diluent is added to the homogeneous solution that reduces the solubility of the product to the point that it “falls out” of solution and crystallizes. It is similar to precipitation but results in the formation of crystals rather than unordered aggregates. Crystallization can be used on a laboratory scale for determining protein structure, on on the industrial scale for antibody and therapeutic protein productions. Batch crystallizers are often used in industry because of their simplicity and inexpensiveness compared to continuous crystallization (Harrison et al., 2003).

==Membrane Separation==
Membrane separation takes advantage of the selective permeability of membranes; they allow certain particles to pass through and selectively stop other, generally unwanted, particles. The component that passes through is called the permeate and the component stream that is rejected is called the retentate or concentrate. The applicability of membranes comes from the fact that their selectivity is determined by their pore size, which can be controlled during the creation of the membranes. Additionally, Membrane processes do not require heat meaning they generally require less energy than conventional separations technology such as distillation and crystallization. Membrane separations processes are generally classified as microfiltration, ultrafiltration, or nanofiltration depending on the size of the particles to be filtered out.

[[File:membrane techs.png|frame|Figure. Cutoffs for different membrane categories]]

===Membrane Selection, Construction, and Flow Geometries===
Membrane permeability and selectivity are the two most important factors to consider when selecting a membrane. For gas separations, the permeation of the gas is usually facilitated by the gas dissolving in the membrane on one side and then evaporating on the permeate side. Therefore permeability depend largely on the solubility of components in the membrane.

The two most commonly utilized membrane configurations are hollow fiber and spiral wound. Hollow fiber is generally the most commonly utilized module for gas separations. These are formed by gluing the two ends of the hollow fiber to a resin forming a closure. The fibers are housed in a shell much like a heat exchanger. The feed flows past thousands of tubes with the permeate flowing into the hollow tubes and out the closure. The retentate then flows out of the shell not having gotten through the membrane.

[[File:hollow membrane.jpg|frame|Figure. Hollow fiber membrane module]]

Spiral wound membranes are created by sealing two membrane sheets back to back on three edges to form a sort of pocket. This fourth open edge is then attached to a porous tube which allows permeate to go through it. Several membrane pockets are attached to a single tube and wrapped around in a spiral.

[[File:spiral membrane.jpg|frame|Figure. Spiral wound membrane module]]

Flow geometry is usually either dead-end geometry or cross flow geometry. In dead end, the fluid flow is normal to the membrane surface while cross flow is parallel to the membrane surface. Dead end geometry is usually used with hollow fiber membranes while cross flow is used with spiral wound membranes. Each geometry has advantages and disadvantages. Dead end geometry is generally cheaper to set up and therefore has lower initial capital costs. However, it is very vulnerable to membrane fouling, which reduces the effectiveness of the membrane. This is usually the geometry set up for small scale lab experiments. The tangential flow devices are more cost and labor-intensive, but they are less susceptible to fouling due to the sweeping effects and high shear rates of the passing flow. Most commercial industrial membrane separations are done using spiral wound cross flow membrane modules.

===Applications===
====Food Industry====
Due to the fact that MD can be conducted at relatively low feed temperatures, it was successfully tested in many areas where high temperature applications lead to degradation of the process fluids especially in food processing. It was demonstrated that MD can be used for the concentration of milk, for the recovery of volatile aroma compounds from black currant juice, and for the concentration of many other types of juices including orange juice, mandarin juice, apple juice, sugarcane juice, etc.
====Reverse Osmosis====
Reverse osmosis is the most widely used membrane separation process. In this process, fresh water passes through the membrane while dissolved salts and other solids are rejected and stay in the concentrate. In this process, feed water is pressurized in order to overcome the osmotic potential difference between the salty retentate and the fresh water desired. These processes are generally run using spiral wound membrane cylinders using a cross flow setup.

===Membrane Model===
The two most important components when considering different membranes are the permeability, which will determine flux through the membrane, and selectivity, which will determine what passes through the membrane and how much. The flux through a membrane is defined as:
<math> M_i = \frac{P_i}{δ}(p_{i,f} - p_{i,p})</math>
Where Mi is the molar flux of component i, Pi is the permeability of the membrane for component i, δ is the membrane thickness, and pi,f and pi,p are the partial pressures of component i on the feed side and permeate side respectively.
The average flux across a long cylindrical membrane such as the spiral wound module is given by:
<math> \int_0^Lm \frac{M_i,dx}{L_m}</math>
Where Lm is the length of the cylinder and x is length in meters

Membrane selectivity of the ideal separation factor is given as the ratio of the permeability of one substance over another as shown:

<math> S_(i,j) = P_i/P_j </math>

where Sij is the selectivity of the membrane for component i over j.

==Cyclones==
Centrifugal Separators, more commonly know as Cyclones, are one of the most widely used gas-solid separators. They are typically used for particles that are 5μm or larger, but if agglomeration occurs they can sometimes separate particles as small as particles as small as .5μm. These cyclones can be designed for high efficiency separations or for high throughput (Towler, 2012).

[[File:cyclone.png|thumb|center|Reverse-flow Cyclone (Towler, 2008)|500x300px]]

===Design===
Design Procedure
There are multiple design methods for cyclones. The deign method discussed below is the Stairmand method. This involved creating two designs one of which was for high throughput and the other which was designed for high efficiency. Using these base designs future cyclones can simply be estimated using scaling factors (Towler, 2012).

1. Select whether efficiency or performance is more important for your design

2. Determine the particle size distribution in your stream

3. Estimate the number of cyclones needed

3. Calculate the cyclone diameter

4. Calculate the scale up factor

5. Calculate the cyclone performance and overall efficiency

6. Calculate the cyclone pressure drop

7. Cost the system

====Particle Size====
The diameter of particles separated by cyclones are governed by this design equation. This equation is dependent upon the standards of the cyclone and how much you are trying to deviate from its standard operating conditions. Looking into the equation it shows that a larger internal cyclone diameter, higher flowrate, higher difference in density, and lower viscosity result in smaller particles being seperated. The whole term after d1 is known as a scale up factor. When designing a cyclone you would determine the flow rate, viscosity, particle size, and density difference would be known prior to using this design equation, so this would be used to optimize the diameter of the cyclone you were designing (Towler, 2012).

[[File:Towler 10.8.PNG|center|Reverse-flow Cyclone (Towler, 2012)|500x300px]]

d1=mean diameter of particle separated at the standard conditions, at the chosen
separating efficiency

d2= mean diameter of the particle separated in the proposed design, at the same
separating efficiency

Dc1 =diameter of the standard cyclone

Dc2 =diameter of proposed cyclone

Q1 =standard flow rate

Q2 =proposed flow rate

Δρ1 =solid-fluid density difference in standard conditions

Δρ2 =density difference, proposed design

μ 1=test fluid viscosity

μ 2=viscosity, proposed fluid.

[[File:Towler10.44.PNG||thumb|center|(a) Standard Cyclone Dimensions (b) High Gas Rate Cyclone (Towler, 2012)|500x300px]]

====Efficiency====
Efficiency is the percentage of total solids that are removed from the stream The graphs below were experimentally for the two standard designs Stairmand Method. This graph has the efficiency vs the particle size.

[[File:Towler10.45.PNG||thumb|center|Performance curves, standard conditions. (a) High-efficiency cyclone performance curves, standard conditions. (b) High gas-rate cyclone (Towler, 2008)|600x400px]]

If you have to scale the efficiency it can be estimated by the point where a horizontal line from the efficiency of the original particle size intersects the new particle size. A demonstration is on the graph below.

[[File:Towler 10.46.PNG||thumb|center|Scaled Performance Curve(Towler, 2008)|500x300px]]

====Pressure Drop====
One of the most important design factors in a cyclone is its pressure drop. The pressure drop occurs because entry and exit losses, kinetic energy losses, and friction losses that occur inside the cyclone. The pressure drop can be calculated using the following equation.

[[File:Towler 10.9.PNG|center|Reverse-flow Cyclone (Towler, 2008)|500x300px]]

The variables are defined as follows:
DP = cyclone pressure drop, millibars;
rf = gas density, kg/m3
;
u1 = inlet duct velocity

u2 = exit duct velocity

rt = radius of circle to which the center line of the inlet is tangential

re = radius of exit pipe

φ =factor from figure below

ψ = fc(As/A1)

fc = friction factor

As = surface area of cyclone exposed to the spinning fluid

A1 = area of inlet duct

[[File:Towler_10.47.PNG|thumb|center|Cyclone Pressure Drop Factor (Towler, 2008)|600x400px]]

====Cost Estimation====
The following values are represented in 2002 dollars. Cost is mostly a function of the amount of gas that need to be passed through the cyclone and the units are in standard meters cubed per second (sm3/sec).

The costs can be broken down into three categories:

Capital Costs: $4600-$7600 sm3/sec

Operations and Maintenance: $1500-$18,000 sm3/sec

Annualized Cost: $2800-$29000 sm3/sec

This all comes to $.47-$440 per metric ton of pollutant removed. The high variation in these costs have to do with smaller units being much less efficient per unit of pollutant removed. Furthermore the cost is also effected by the amount of particulates that are in the gas to begin with. For example a gas with 10 g/sm3 and air with 100sm g/sm3 would have similar pumping costs, but very different cost per unit of solids removed.

==Other Separation Processes==
===Extraction===
Liquid-liquid extraction is a process for components with overlapping boiling points and azeotropes. The process requires a solvent such that some of the components of the mixture are soluble, and then the components will be separated based on this solubility in the liquid. This process can operate at moderate temperatures and pressures, so is not very energy intensive. However, a distillation column is required to extract the solvent for recycle. More recently, supercritical fluids have replaced liquid solvents in some processes for L/L extraction, due to the solute’s ability to more rapidly diffuse through them. The issue with these fluids, however, is that they must be operated at extremely high pressures and temperatures, increasing both capital and operating expenses of the process (Peters & Timmerhaus, 2003).

===Crystallization===
This process recovers solutes that have been dissolved in solution. The resulting product is in the solid phase. Depending on the material properties of the solute and solvent, the solute is recovered by precipitation after cooling, removal of solvent, or adding precipitating agents. Crystallizers are designed based on phase equilibria, solubilities, rates and amounts of nuclei generated, and rates of crystal growth. Every crystallization process is a unique system, so plant evaluation is usually required before complete implementation. Crystallization can be performed in both batch and continuous processes, and design features can control crystal size to an extent (Peters & Timmerhaus, 2003).

===Membrane Separation===
This separations process uses selectively permeable membranes to separate components in a mixture. Typically, one of the components will freely pass through the barrier while the other components will not. The stream that passes through the membrane is the permeate and the stream that does not pass is the retentate. The driving force behind this separation is a pressure gradient. Membrane separation is beneficial because it can separate mixtures at the molecular and small particle level. Furthermore, there is no phase change required so the energy input is low. Limitations of this process include achieving high product purity, incompatibility with certain stream components, low operating temperature, and low flow rates. Although membrane separation is generally not scaled up, examples of scaled-up membrane separation include seawater desalination and hydrogen recovery (Peters & Timmerhaus, 2003).

===Adsorption===
Adsorption involves an adsorbent and adsorbate. The adsorbent is typically a solid, and will typically separate the adsorbate from the stream. This process usually includes a desorption step that regenerates the adsorbent for further use. Raising the temperature or increasing the concentration of the adsorbate can reverse the adsorption process. Although the recycle of the adsorbent is a very economic design feature, the downside of this step is that it results in a cyclic process, which introduces complexity to the overall process. Industrial applications of this process are for bulk separations and gas purification. The adsorption/desorption process in these situations involves a large amount of heat transfer, which design engineers must take into account when sizing and selecting equipment material (Peters & Timmerhaus, 2003).

===External Field/Gradient Separation===
These separations use external force fields or temperature gradients to separate responsive molecules or ions. The use of these processes is fairly limited to a few specialized industrial applications (Peters & Timmerhaus, 2003).

===Settling and Sedimentation===
In settling processes, solid particles or liquid drops are separated from a stream by gravity. The stream can be in either the liquid or gas phase. For vapor-liquid mixtures, flash drums are generally used to separate the mixture. The velocity of the vapor must be less than the settling velocity of the liquid drops for this separation to occur. For liquid-liquid separation, the horizontal velocity of the fluid must be low enough to allow the low-density droplets to rise to the interface and the high-density droplets to move away from the interface and coalesce. In sedimentation, the result of the process is a more concentrated slurry. Typically a flocculating agent is used to aid in the settling process. One way to perform this separation is to use a cone-shaped tank with a slowly revolving rake that scrapes and moves the thickened slurry to the center of the cone for removal (Peters & Timmerhaus, 2003).

[[File:Example.jpg]]

====Clarifiers====
[[File:Circular_Clarifier.png|300px|thumb|bottom|Figure 9: Circular clarifier with some components labelled.]] [[File:Rectangular_Clarifier.png|300px|thumb|bottom|Figure 10: Rectangular clarifier with some components labelled.]]

Clarifiers are one of the methods used for the continuous removal of particulate solids from liquids through sedimentation by gravity. Applications include process water pretreatment, waste water treatment, and drinking water purification. Historically, clarifiers were originally developed to limit nutrient input into surface water due to fear of eutrophication. Today, they have a number of uses, particularly in wastewater treatment processes, metal removal, disinfection, and membrane pretreatment. The process helps removed dissolved solids, silt, and undesirable metals from the water, making it more suitable for downstream processes as well as human consumption (Wilson, 2005).

Clarifiers are typically used in conjunction with coagulation or flocculation agents, which promote dissolved particles to join into clumps and settle out of solution (Towler and Sinnot, 2012). Clarifiers typically consist of a large circular tank with a rotating rake at the base which scrapes settled solids towards the center. In the case of a rectangular clarifier, they are scraped to one side. Diagrams of both are represented in figures 9 and 10, respectively (NMED Surface Water Quality Bureau, 2015). Separated solids are allowed to settle to the bottom of the tank as a sludge, whereupon they are collected by the rake and disposed of properly. In the case of floating contaminants, it is possible for the clarifier to include a skimmer as well.

Clarifier efficiency varies with certain factors, including the settling characteristics of solids removed and the surface overflow rate of the tank. Clarifier efficiency can be found using the following relation:

<math>\begin{align}
E_{TSS} &= E_{TSSmax}\left ( 1 - e^\frac{\lambda}{SOR} \right )
\end{align}</math>

where <math>E_{TSS}</math> is the efficiency of total suspended solids (TSS) removal, <math>E_{TSSmax}</math> is the maximum possible efficiency, <math>\lambda \left [\frac{m}{d} \right ]</math> is the settling constant, and <math>SOR \left [\frac{m^3}{m^2 d} \right ]</math> is the surface overflow rate. The effect of flocculation chemicals on TSS can be seen in figure 11. However, it should be noted that chemical addition will increase sludge quantity and may have an adverse effect on plant aesthetics, which increases maintenance costs (Wilson, 2005).

[[File:Chem_Addition.png|200px|thumb|bottom|Figure 11: The effect of flocculating agents on total suspended solids removal in clarifiers.]]

=====Lamella Clarifiers=====

Lamella clarifiers use inclined plates in order to maximize the settling area for solids. Solids continue to settle into a hopper at the bottom of the tank while clarified water exits up through the inclined plates. This allows for the design of a smaller tank, which leads to large savings in capital costs. A lamella clarifier is pictured in figure 12.

[[File:Lamella_Clarifier.png|300px|thumb|bottom|Figure 12: A lamella clarifier with components labeled.]]

Typically, inclined plates are installed at an angle of 45 to 60 degrees and spaced 40 to 120 mm apart, which increases effective settling surface area by a factor of 6 to 12 compared to traditional clarifiers. For effective use, it is recommended that the Reynolds number be below 2000, Froude number higher than 10-5,and detention time be longer than 3 to 5 minutes. For this implementation, the equations are as follows:

<math>\begin{align}
N_{Re} &= \frac{VR}{\nu} \\
N_{Fr} &= \frac{V^2}{Rg}
\end{align}</math>

where <math>R</math> refers to the hydraulic radius, which is the cross-sectional area of the lamella, <math>V</math> is the liquid velocity, <math>\nu</math> is the kinematic viscosity, and <math>g</math> is the gravitational constant (Wilson, 2005).

=====Advantages=====

Clarifiers offer a proven, relatively inexpensive solution for solids removal. The chemical coagulants used are cheap and provide a low operating cost as well as simple maintenance. Construction is typically simple, leading to low capital costs and equipment that is easy to accommodate and maintain. Their design is also flexible, with various options such as skimmers and scrapers offering increased removal efficiency (Wilson, 2005). Operation of clarifier tanks also has lower energy requirements than membrane filtration for solids removal, given that most of the separation is aided by gravity. Water exiting clarifier units has a silt density index (SDI) averaging 4.0, which is low enough for further membrane treatment such as reverse osmosis (Prihasto, 2009).

=====Disadvantages=====

Clarifiers necessitate low turbulence to prevent resuspension of solids. This essentially requires a low entrance velocity, which can limit the production rate of certain processes or call for more clarifier units, which would drive up costs. Furthermore, clarifiers require frequent cleaning before sludge becomes too difficult to remove and reduces effectiveness. In the case of lamella clarifiers, sludge buildup on the inclined plates results in uneven flow distribution which could harm efficiency (US EPA, 2003). For this reason, maintenance requirements for lamella clarifiers are higher, but they can be reduced through the implementation of removable plates (Wilson, 2005). Clarifiers also only remove solids, so pH will not be affected, leading to the need for further pH adjustment (NMED Surface Water Quality Bureau, 2015).

=====Clarifier Design Calculations and Typical Design Values=====

======Detention Time======

Detention time (DT) is the time is takes for a unit of water to travel from the inlet of the clarifier unit to the outlet. During typical operations, the design value for this is 2 to 3 hours.

<math>\begin{align}
DT &= \frac{Tank\ Volume}{Influent\ Rate}
\end{align}</math>

======Surface Overflow Rate======

Surface overflow rate (SOR) measures the flow into the clarifier per square foot of surface area. Typical design values are 400 to 800 gal/day/sq. ft.

<math>\begin{align}
SOR &= \frac{Volumetric\ Flow\ Rate}{Surface\ Area}
\end{align}</math>

======Weir Overflow Rate======

Weir overflow rate (WOR) describes the flow in gallons per day per linear foot of weir. Typical values are 10,000 gal/day/ft.

<math>\begin{align}
WOR &= \frac{Volumetric\ Flow\ Rate}{Weir\ Length}
\end{align}</math>

======Solids Loading Rate======

Solids loading rate (SLR) describes the mass of solids in the clarifier influent per square foot of surface area. This value should not exceed 30 lbs/day/sq. ft.

<math>\begin{align}
SLR &= \frac{Solids\ Mass\ Flow\ Rate}{Surface\ Area}
\end{align}</math>

===Flotation===
Flotation is a process designed for specific solid-solid mixtures. It works by generating gas bubbles in a liquid that attach to selected solid particle. Afterwards, the particles rise to the liquid surface where they are removed by an overflow weir or mechanical scraper. The separation depends on the surface properties of the particles and its preference to attach to the gas bubbles. To meet the necessary requirements of the flotation process, a number of additives can be used to control things like the pH of the liquid-solid mixture, the activity of the solid surface, and the froth that can assist in separation. The bubbles can be produced by gaseous dispersion, dissolution, or electrolysis of the liquid (Peters & Timmerhaus, 2003).

===Centrifugation===
This process is similar to external field separation in that an external force field is applied to separate a mixture. When gravity separation is too slow due to particle densities, particle size, settling velocity, or the formation of an emulsion, centrifugation is commonly used. Centrifugal force increases the total force acting on the particle and results in faster separation times. This process is generally used to separate solids from liquids, however it can also be used to separate two liquids with very different densities (Peters & Timmerhaus, 2003).

===Drying===
Drying is performed to remove liquid from a liquid-solid mixture and produce a dry solid. Water is most often the liquid removed, but organic liquids are removed from solids on occasion as well. The heat required to vaporize the liquid is usually obtained by a series of gas-solid contacting devices. Feed condition and temperature sensitivity of the solid dictate the type of contacting device that is used. There are two groups of dryers that differ by the dependence of either mechanical means or fluid motion for gas solid contact. Another feature of dryers is to use either direct (hot gas) or indirect (conductive surface) heating (Peters & Timmerhaus, 2003).

===Evaporation===
Evaporators separate solvents from a solution by evaporation. The difference between evaporation and distillation is that evaporation requires the solute be nonvolatile. Because of this, a high separation can be achieved with one stage. Evaporators are essentially reboilers, so evaporation is a very energy-intensive process with a high thermal economy (Peters & Timmerhaus, 2003).

===Filtration===
Filtration is a process that separates a mixture of solid in a liquid or gas by passing the mixture through a porous medium in which the particles do not pass. Filtration is done by either cake filtration (particles found on the surface of the filter) or depth filtration (particles found within the filter). Cake filtration is generally performed with a cloth as the filtration medium (Peters & Timmerhaus, 2003).

==Conclusion==
Separation is a key part of most chemical processes, and there is a great variety of techniques to perform separation of compounds based on size, volatility, charge, and many other features. A common technique with which the process engineer should be familiar is distillation, but he or she should also be aware of the other available options. Some techniques may be less expensive, less energy-intensive, or more effective than distillation, depending on the specific separation problem. Therefore, the separation strategy should be carefully considered.

==References==
Air Control Technology Fact Sheet. Enviornmental Technology Fact Sheet. http://www3.epa.gov/ttncatc1/dir1/fcyclon.pdf

Belter PA, Cussler EL, Hu WS. Bioseparations: Downstream Processing for BIotechnology. New York: John Wiley; 1998.

Biegler LT, Grossmann IE, Westerberg AW. Systematic Methods of Chemical Process Design. Upper Saddle River: Prentice Hall; 1997.

Danckwerts P (1965) The Absorption of Gases in Liquids. Pure and Applied Chemistry UK 10:625-642.

Development Document for the Final Effluent Limitations Guidelines and Standards for the Metal Products and Machinery Point Source Category (Report). US Environmental Protection Agency. 2003.

Erwin, D. Industrial Chemical Process Design. New York: McGraw Hill, Professional Engineering; 2002.

Hage, David S. 1999Affinity Chromatography: A Review of Clinical Applications. Clinical Chemistry 45(5): 593–615.

Harrison RG, Todd P, Rudge SR, Petrides, DP. Bioseparations Science and Engineering. New York: Oxford University Press; 2003.

Hage, David S. 1999Affinity Chromatography: A Review of Clinical Applications. Clinical Chemistry 45(5): 593–615.

HisPur Cobalt Resin - Thermo Fisher Scientific N.d. https://www.thermofisher.com/order/catalog/product/89964, accessed February 21, 2016.

Lamella Plate Clarifier. Hydro International Web site. Available at: http://www.hydro-int.com/uk/products/lamella-plate-clarifier?s=0&r=uk. Accessed February 2, 2016.

Lean Oil Absorption. PetroGas Systems Web site. Available at: http://petrogassystems.com/technology/natural-gas-processing-and-dew-point-control/lean-oil-absorption. Accessed February 19, 2014.

Lichty, Jordan J., Joshua L. Malecki, Heather D. Agnew, Daniel J. Michelson-Horowitz, and Song Tan 2005Comparison of Affinity Tags for Protein Purification. Protein Expression and Purification 41(1): 98–105.

Merichem Gas Technologies. ®LO-CAT PROCESS available at http://www.merichem.com/images/casestudies/Desulfurization.pdf Accessed 6 Feb. 2015.

Miller L.N. & Zawacki T.S. , US 4080424, "Process for acid gas removal from gaseous mixtures", issued 21 Mar 1978, assigned to Institute of Gas Technology

NMED Surface Water Quality Bureau, New Mexico Water Systems Operator Certification Study Manual, New Mexico Environment Department, 2015.

Peters MS, Timmerhaus KD. Plant Design and Economics for Chemical Engineers. 5th ed. New York: McGraw Hill; 2003.

Prihasto, N; Lui, Q; Kim, S. Pre-treatment strategies for seawater desalination by reverse osmosis system. 2009; 249(1): 308-316. doi:10.1016/j.desal.2008.09.010

Schmidt Eberhard (2012) Waste Gases, Separation and Purification. Ullman’s Encyclopedia of Industrial Chemistry Germany 2:174-181.

Seider, W.D., Seader, J.D., and Lewin, D.R. (2004). ''Process Design Principles: Synthesis, Analysis, and Evaluation.'' New York: Wiley.

Stripping Column. Alfa Laval Web site. Available at: http://www.alfalaval.com/solution-finder/products/soft-column/Documents/Stripping%20Column.pdf. Accessed February 19, 2014.

Towler G, Sinnott R. Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013.

Turton, R.T., Bailie, R.C., Whiting, W.B., and Shaewitz, J.A. (2003). ''Analysis, Synthesis, and Design of Chemical Processes'' Upper Saddle River: Prentice-Hall.

Wankat, P.C. (2012). ''Separation Process Engineering.'' Upper Saddle River: Prentice-Hall.

Wilson, T.E., Clarifier Design, 2nd Ed., McGraw-Hill: New York, 2005.

Separation processes

2016-02-22T01:26:35Z

Bts511: /* Cyclones */

Authors: Nick Pinkerton, [2014] Karen Schmidt, [2014] James Xamplas, [2014] Emm Fulk, [2015] and Erik Zuehlke, [2015] John Dombrowski [2016] , Brett Sleyster [2016] , Robert Cignoni [2016] , and Osman Jamil [2016] 

Stewards: David Chen, Jian Gong, and Fengqi You

Date Presented: February 9, 2014 /Date Revised: February 1, 2014

 

==Introduction==
Essentially all chemical processes require the presence of a separation stage. Most chemical plants comprise of a reactor surrounded by many separators. Separators have a countless number of jobs inside of a chemical plant. A separator can process raw materials prior to the reaction, remove incondensable gases, remove undesired side products, purify a product stream, recycle materials back into the process, and many other jobs that are essential to the process.

Chemical engineers must understand the science of separation and the variety of ways that separation can take place. There are many ways to perform a separation some of these including: distillation, absorption, stripping, and extraction. The science of separation revolves around the presence of two phases that are in contact and equilibrium (Wankat, 2012).

[[File:Sepmeth.JPG|frame|Figure 1. Separation methods by property]]

==Theory==
===Vapor-Liquid Equilibrium===
Separation processes are based on the theory of vapor-liquid equilibrium. This theory states that streams leaving a stage in a separation process are in equilibrium with one another. The idea of equilibrium revolves around the idea that when there is vapor and liquid in contact with one another they are in constantly vaporizing and condensing. Different components in the mixture will condense and vaporize at different rates. There are three types of equilibrium conditions that can be subdivided into thermal, mechanical and chemical potential categories. These separate equilibrium states are given as:

<math>T_{liquid} = T_{vapor}</math>

<math>p_{liquid} = p_{vapor}</math>

<math>chemical potential_{liquid} = chemical potential_{vapor}</math>

==Distillation==
===Flash Distillation===
Flash Distillation is one of the simpler separation processes to be employed in a chemical plant. The main premise of flash distillation is that a portion of a liquid feed stream vaporizes in a flash chamber or a vapor feed condenses. Vapor-liquid equilibrium will cause the vapor phase and the liquid phase to have different compositions. The more volatile component of the mixture will compose of a larger portion of the vapor. This simple separation is easy to manufacture but does not result in large degrees of separation.

Flash distillation requires a feed stream that is pressurized and heated and then passed through a valve into a flash drum. The large pressure drop across the valve will result in a partial vaporization of the fluid. Vapor will be removed overhead from the flash drum while the remaining liquid will collect at the bottom of the drum and be removed. Most flash drums will contain an entrainment eliminator which is a screen that prevents liquid from being carried into the vapor effluent. Figure 2 shows a simple overview of the flash distillation process. As shown, there is a heater that flows into a let-down valve where the two-phase flow begins. Variables y and x are the mole fractions of the more volatile component in the vapor and liquid effluents, respectively.

[[File:Flash.gif|center|frame|Figure 2. Flash Distillation Flow Diagram]]

===Column Distillation===
Distillation columns are the most widely used separation technique used in the chemical industry, accounting for approximately 90% of all separations (Wankat, 2012). Distillations in columns consist of multiple trays that each act at their own equilibrium conditions. Large columns are able to perform complete separations of binary mixtures as well as more complex multi-component mixtures.

[[File:column.jpg|250px|center|]]
===Stages===
Columns are separated into stages by the presence of trays. These trays allow for vapor-liquid contact and equilibrium to occur. Typically, the more stages in a column, the larger separation that can be achieved. There are many different types of trays that can be used in a column.
====Sieve Trays====
The simplest and least expensive tray type is the sieve tray which is a sheet of metal with holes punched into it to allow vapor flow. Sieve trays can have different hole patterns and sizes that will affect the tray efficiency and flow rates.

[[File:sieve.jpg|200px|center|]]

====Sieve Tray Design Procedure====

The design of these plates is done through a trial-and-error process. Most commercial process simulations (such as HYSYS) have default tray designs, and automatically specify dimensions. However, these dimensions selected or calculated by the simulations may not give the best performance for your system, so it is valuable to understand how to design the sieve trays and how specific parameters may affect performance. Hand calculations using the following methods can be used to guide the simulation programs to better design. This section will use sample data to work through an example of the process. The following is a general list of steps for designing a sieve plate:

=====1. Calculate the maximum and minimum vapor and liquid flow rates for the turndown ratio required.=====
This data can be collected from a McCabe-Thiele diagram and/or from process simulation data.

Data from McCabe Thiele diagram, for example: 
Number of stages = 10 
Slope of top operating line = 0.185 
Slope of bottom operating line = 1.43 
Top composition = 98.8 mol% acetone 
Bottom composition = 4 mol% acetone (a) 
Minimum reflux ratio = 0.31 

Collect data rom mass balances: 
MW of feed = (0.6 kmol a/kmol)*(58 kg/kmol a) + (0.4 kmol acetic acid/kmol)*(60 kg/kmol aa) = 59 kg/kmol 
Feed rate = 100 kmol/h 
Using material balances... D = 57.3 kmol/h ; B = 42.7 kmol/h 
Vapor rate, V = D(1+R) = 57.3 kmol/h (1+.31) = 75.1 kmol/h

=====2. Collect or estimate the system physical properties.=====
Here it is important to know information about both the top and bottom of the column. Useful information includes temperature, pressure, column pressure drop (a common assumption is 100 mmH2O per plate), densities, molecular weights, surface tensions, and number of stages (which can be estimated from the McCabe-Thiele diagram).

=====3. Select a Trial Plate Spacing=====
The plate spacing will depend on the column diameter and operating conditions. Plate spacings from 0.15 m to 1.0 m are typically used. The smaller the diameter, the smaller the spacing. Small columns will use close spacing. Columns with diameters above 1.0 m, plate spacings of 0.3 m to 0.6 m are normally used. A good initial estimate is 0.5 m.
[[File:trayspacing.jpg|400px|center|]]

=====4. Estimate the column diameter, based on flooding considerations.=====
Vapor and liquid flow rates will vary along the column, so plate design needs to be considered both above and below the feed. Using plate spacing and FLV (which is the square root of the ratio of the liquid to vapor flow rates), you can obtain the value of K from the plot.

[[File:floodingplot.jpg|400px|center|frame|Figure. Plate Spacing]]
There is a range of vapor and liquid flow rates in which the column needs to be operated. Too low or too high of rates can result in various inefficiencies in the column operation, as shown in the figure below. For example, if the vapor rate is too high, flooding will occur. However, it is not safe to operate on the flooding line. Instead, columns are typically designed for 80% of flooding at the maximum flow rate.
[[File:vap_rate_vs_liq_rate.jpg|400px|center|frame|Figure. Tray behavior]]

=====5. Decide the liquid flow arrangement.=====
Common flow arrangements are single pass (cross flow), double pass, and reverse flow. Using conditions at the bottom of the column, calculate the max volumetric flow rate. Use this flow rate and the column diameter to determine the preferred flow arrangement from the chart below.
[[File:Liquidflow.jpg|200px|center|]]

=====6. Make a trial plate layout: downcomer area, active area, hole area, hole size, weir height.=====
Standard sizes for trays -- and good assumptions for the first iteration -- are: weir height, hw = 50mm ; hole diameter, Dh = 5mm ; plate thickness, tpl = 5mm. From the graph below, the ratio of downcomer area (Ad) to column cross-sectional area (Ac) can be determined from the ratio of weir length (lw) to column diameter (Dc) and vice versa.
[[File:platelayout.jpg|200px|center|frame|Figure. Plate Dimensions]]

=====7. Check the weeping rate=====
Compare the actual vapor velocity to the minimum vapor velocity -- if velocity is too low fluid will "weep" through the tray holes. If the weeping rate is unsatisfactory, return to step 6 and choose different values for the plate layout dimensions. From the chart in step 4, it can be seen that there is a minimum vapor flow rate below which the liquid "weeps" from the tray above.

For the remaining steps in this design process, it is recommended to check your assumptions after each step and revise them as necessary in order to maintain operation in the "sweet spot" of the vapor rate vs. liquid rate plot. Additional iterations may be required as you move through the procedure.

Calculate the maximum liquid flow rate. Calculate the minimum liquid flow rate at 70% turndown (recommended). Calculate the height over the weir as
<math>h_o=750[\frac{L_w}{p_Ll_w}]^\frac{2}{3}</math>

=====8. Check the plate pressure drop=====
<dfn>If the pressure drop calculated here is too high, return to step 6.</dfn>

Proceed to step 9 if the pressure drop assumption is valid.
=====9. Check the downcomer backup. =====
<dfn>If the downcomer backup is too high, return to step 6 or 3.</dfn>

The plate spacing affects the amount of fluid in the downcomer. Calculate the level in the downcomer and the residence time of the fluid to see if the values are valid. Note that residence times greater than 3 seconds are acceptable.

Proceed to step 10 if residence time is acceptable.
=====10. Decide plate layout details.=====
Determine calming zones, the unperforated areas at the inlet and outlet sides of the plate. The width of each zone is usually made the same. Recommended values are: below 1.5 m diameter, 75 mm; above, 100 mm. The unperforated area can be calculated from plate geometry. Also check the hole pitch, or the distance between hole centers. It should not be less than 2.0 hole diameters. A normal range is between 2.5 and 4.0 hole diameters. The shape must also be specified. Square and equilateral triangle holes are used.

=====11. Recalculate the percentage flooding based on the chosen column diameter.=====
An assumption of 80% flooding was chosen so that operation would occur in the "sweet spot." This assumption must be checked by calculating the flooding percentage for a given column diameter.
uv = (max volumetric flow rate)/(net area)

<math>%flooding = \frac{u_v}{u_f}</math>

If the hole pitch is unsatisfactory, return to step 6.
=====12. Check entrainment=====
''If too high, return to step 4'' Use the graph below to determine entrainment from FLV.
[[File:entrainment.jpg|400px|center|]]

The value for fractional entrainment can be used to re-estimate the column efficiency, and reevaluate the number of trays needed. Can return to step 1 for more accurate estimates.

=====13. Optimize design.=====
After returning to step 1 to reevaluate the number of trays, it is valuable to repeat steps 2 through 12 to find the smallest diameter and plate spacing acceptable at the lowest cost.

=====14. Finalize the design.=====
Optional: draw up the plate specification and sketch the layout of the plate.

====Bubble-Cap Trays====
Bubble-cap trays consist of a weir around each hole in the tray which is covered with a cap that has holes or slots to allow vapor passage. Entrainment is about three times larger than a sieve tray. Bubble-cap trays require larger tray spacing than sieve tray design. Bubble-cap trays have been known to have problems with coking, polymer formation, or high fouling mixtures. Recently, very few new bubble-cap columns are being built due to the expense and marginal benefits. However, engineers will likely encounter bubble-cap columns still currently in operation.

====Flow Patterns====
Cross flow columns are the most common pattern for distillation columns. For liquid flows between 50 and 500 Gal/min, a cross flow column is appropriate. When liquid flow is increased above 500 Gal/min, an engineer should consider designing a double pass or multi-pass column. This will reduce the liquid gradient on the tray and reduce the downcomer loading (Wankat, 2012).

===Column Sizing===
Column height will be dependent on the amount of trays required and the spacing between the trays. Normally, tray spacing of 0.15 m to 1 m is used. For columns, above 1 meter in diameter, 0.5 m can be used as an initial estimate.

Column diameter is influenced by the vapor flow rate in the column. The trays can not have excess liquid entrainment or high pressure drops; therefore, vapor velocity in the column must be maintained at a reasonable level.

An equation based on the Souders and Brown equation can be used as an estimate for the max allowable superficial vapor velocity,

<math>\hat u_v = (-0.171l_t^2 + 0.27l_t - 0.047){\frac{\rho_L - \rho_v}{\rho_v}}^{1/2}</math>

where <math>l_t</math> is the plate spacing in meters, <math>\rho_L</math> is the density of the liquid stream, and <math>\rho_V</math> is the density of the vapor stream.

Column diameter, <math>D_c</math>, can then be estimated using the relation,

<math>D_c = \sqrt{\frac{4\hat{V_w}}{\pi\rho_v\hat{u_v}}}</math>

where <math>\hat{V_w}</math> is the maximum vapor rate in kg/s (Towler et al., 2013).

===Distillation Applications===

Distillation is a process that can be implemented in various scales. There is both laboratory scaled distillation as well as very large industrial distillation. Other applications for distillation include food/alcohol processing and herb distillation for the perfume and medical industries. Typically laboratory scaled distillation occurs in batches whereas industrial distillation (e.g. fractional distillation of crude oil) occurs continuous with a constant distillate and bottom effluent streams.

Some applications of distillation are concerned the top stream only, some the bottom stream only and others both streams can be used for future products. In alcohol distillation for example, the water that is separated from the ethanol/water binary solution is discarded as waste water. In fractional distillation of crude oils, the heavy hydrocarbons at the bottom of the column are collected and sold along with the light hydrocarbons that appear in higher side draws (Wankat, 2012).

===Example Case: Ideal Distillation===

Assume an equimolar mixture flowing at 10 mol/s of 20 mol% n-pentane, 30 mol% n-hexane, and 50 mol% n-heptane. Separate the mixture into 3 products: 99% pure n-pentane, 99% pure n-hexane, 99% n-heptane. Assume the feed and products are all liquids at the bubble points. There are two process alternatives to consider in this example. The direct sequence removes the most volatile species, pentane, in the first column, and then separates hexane and heptane in the second column. The indirect sequence separates the heaviest product, heptane, and then separates pentane from hexane in the second column. This example will consider the direct sequence. Next, we must decide if these species exhibit fairly ideal behavior during distillation. Since the n-alkanes have very similar properties, it is safe to assume they will display close to ideal behavior. The next step is to look up the boiling points of the 3 species. In this case, the normal boiling points of pentane, hexane, and heptane are 309 K, 342 K, and 372 K, respectively. Also, it is a good idea to look up relative volatilites, to further verify near-ideality of the mixture, but also to obtain the information necessary for the Underwood method, which we will employ to obtain a solution. The next step is to write out material balances based on molar flows and the design specifications. They go as follows:

<math>\mu_I(nC5) + \mu_{II}(nC5) = 2 mol/s</math>

<math>\mu_I(nC6) + \mu_{II}(nC6) + \mu_{III}(nC6) = 3 mol/s</math>

<math>\mu_{II}(nC7) + \mu_{III}(nC7) = 5 mol/s</math>

<math>\mu_I(nC5) = 99\mu_I(nC6)</math>

<math>\mu_{II}(nC5) = (5/990)\mu_{II}(nC6)</math>

<math>\mu_{II}(nC7) = (5/990)\mu_{II}(nC6)</math>

<math>\mu_{III}(nC7) = 99\mu_{III}(nC7)</math>

where <math>\mu</math> represents the molar flow, and the subscript represents the product stream.

Solving this system of equations:

<math>\mu_I(nC5) = 1.985\ mol/s</math>

<math>\mu_{II}(nC5) = 0.015\ mol/s</math>

<math>\mu_I(nC6) = 0.020\ mol/s</math>

<math>\mu_{II}(nC6) = 2.930\ mol/s</math>

<math>\mu_{III}(nC6) = 0.050\ mol/s</math>

<math>\mu_{II}(nC7) = 0.015\ mol/s</math>

<math>\mu_{III}(nC7) = 4.985\ mol/s</math>

At this point we have enough information to use Underwood's method to estimate the minimum vapor flows in the column. The following three equations are used in Underwood's method:

<math>\sum_i \frac{\alpha_{ik}}{\alpha_{ik}-\phi}f_i = (1-q)F</math>

<math>(R_{min}+1)D = \sum_i \frac{\alpha_{ik}}{\alpha_{ik}-\phi}d_i = V_{min}</math>

<math>\bar R_{min}B = -\sum_i \frac{\alpha_{ik}}{\alpha_{ik}-\phi}b_i = \bar V_{min}</math>

where <math>\alpha_{ik}</math> is the relative volatility of species <math>i</math> to species <math>k</math>, <math>f_i</math> the molar flow of species <math>i</math> in the feed, <math>q</math> the fraction of the feed that joins the liquid stream at the feed tray, <math>F</math> the total molar flow of the feed, <math>D</math> the molar flow of the distillate, <math>R_{min}</math> the minimum reflux ratio <math>(=L_{min}/D)</math>, <math>d_i</math> the molar flow of species <math>i</math> in the distillate, <math>V_{min}</math> the minimum vapor flow possible in the top section of the column to accomplish the desired separation, <math>\bar R_{min}</math> the minimum reboil ratio <math>(=\bar V_{min}/B)</math>, <math>b_i</math> the molar flow of species <math>i</math> in the bottoms product, and <math>\bar V_{min}</math> the minimum vapor flow in the bottom section of the column. The final variable, <math>\phi</math>, will be solved for using the first Underwood equation, and it's value will be decided based on the relative volatilities of the key components in the column.

So, after solving the first Underwood equation, we get two values for <math>\phi</math>, 3.806 and 1.462. Because 3.806 is between the relative volatilities of the key components, we will substitute that value for <math>\phi</math> into the second Underwood equation. Doing so for both columns gives <math>V_{min} = 6.4\ mol/s</math> for the first column and <math>V_{min} = 8.9\ mol/s</math> for the second column, for a total minimum vapor flow of 15.3 mol/s. The process would then be repeated for the indirect sequence, and the decision for which process to use would be justified by the process with the overall minimum vapor flow (Biegler et al., 1997).

==Absorption==
===Description of Absorption===
Another separation process used in industry is absorption, which is used to remove a solute from a gas stream. It accomplishes this by contacting the gas mixture with a liquid solvent that readily absorbs the undesirable components from the gas stream, purifying the gas stream. This separation process is determined by the inputs of the liquid flow rate, temperature, and pressure.

The absorption factor, which can be determined mathematically, determines how readily a component will absorb in the liquid phase. The absorption factor of component i is

<math>A_i=L/K_iV</math>

where <math>L</math> is the liquid flow rate entering the column, <math>V</math> is the vapor flow rate entering the column, and <math>K_i</math> is the vapor/liquid equilibrium ratio for component i (Peters & Timmerhaus, 2003). Higher absorption factors result in higher absorptivity into the liquid and a decrease in the number of trays required for separation, however a diminishing return occurs after the absorption factor is greater than 2.0. An absorption factor of 1.4 is most commonly used.

In general absorption can be seperated into two overarching categories, physical and chemical absorption. In physical absorption, the unwanted solute in the gas is absorbed into the liquid phase because solubility of the component is higher in the liquid phase than the gas phase. In chemical absorption the solute is removed from the gas via a reaction with the solvent, this reacted product is then transported into the liquid phase (Danckwerts 1965). There are two types of chemical absorption reversible and irreversible. Generally reversible chemical absorption is preferred as the solvent can be put through a stripper and regenerated so it can be recycled back to the absorption process (Wankat, 2012).

===Absorption Apparatus===

There are five major apparatus used for absorption in industrial application. These five pieces of equipment are spray absorbers (or towers), ejector (venturi) scrubbers, packed columns, trayed columns, and film absorbers (Schmidt, 2012).

==== Spray Tower vs Ejector Scrubber ====

In both '''spray tower''' and the '''ejector scrubber''' nozzles are employed to produce small solvent droplets. These small droplets increase the surface area of the liquid to gas contact allowing for the maximum amount of mass transfer to occur between the gas mixture and the liquid. The major difference between the two nozzle equipment designs is the configuration and type of nozzles. In the ejector scrubber shown in Figure 3 there is a single nozzle that is generally a higher pressure spray nozzle that produces finer solvent drops allowing for an even greater amount of mass transfer enabling better physical absorption (Schmidt, 2012).
[[File:Ejectorventuri.jpg|thumb|200px|center|Figure 3. Ejector Scrubber (US EPA, 2006)]]
'''Spray towers''' on the other hand generally have many nozzle at different heights where the liquid solvent will be sprayed out of to contact the gas running through the tower. This design is used in order to ensure the gas contacts the liquid as throughout the tower. These nozzles are lower pressure than a ejector scrubbers nozzle and thus physical mixing is worse in this configuration. Since physical mixing is generally worse in this configuration it is usually used in conjunction with a chemical absorption process. The other major difference between the ejector scrubber and the spray tower is that gas and liquid flow is cocurrent in the former while it is countercurrent in a spray tower. A spray tower absorber is shown below in Figure 4 (Schmidt, 2012).
[[File:SparyTowerAbsorber.jpg|thumb|200px|center|Figure 4. Spray Tower Absorber (US EPA, 2006)]]

==== Tower Type Absorption Apparatus ====
'''Packed column absorbers''' and '''tray column absorbers''' have very high efficiencies for the removal of an unwanted solute in the gas stream. The major disadvantage a trayed column has when compared to a packed column is the pressure drop. The pressure drop in a packed column is generally very low, whereas in between each tray of a trayed column pressure drop can be quite large. However the advantages inherent to trayed columns become clear when one needs the solvent to have a high concentration of the component to be removed from the gas stream. This is most important in the case where there is a very low concentration of the component in the gas stream and the specification states the solvent must contain a high concentration of that component. In this case the flow rate of the solvent may not be high enough for a packed column, however in a trayed column the solvent flow rate can be near zero for operation (Schmidt, 2012). Packed and trayed column internals are very similar to the setups found in the respective distillation columns.

For a '''trayed column''' the plate efficiency can be calculated using O'Connell's Correlation which invovles the Henry's Law constant, total system pressure, and solvent viscosity at the operating temperature (Towler & Sinnott, 2013).

<math>x=0.062*\frac{\rho_s*P}{\mu_s*H*M_s}</math>

where
<math>x</math> is the tray efficiency,
<math>\rho_s</math> is the density of the solvent in <math>kg/m^3</math>,
<math>P</math> is the total pressure of the system in <math>N/m^2</math>,
<math>\mu_s</math> is the solvent's viscosity in <math>mNs/m^2</math>,
<math>H</math> is the Henry Law constant in <math>1/(Nm^2*(mol fraction))</math>,
and <math>M_s</math> is the molecular weight of the solvent.

A packed towers height can be determined using the equations below when concentration of solute is below 10% so that the assumption that the flow of gas and liquid will be essentially constant throughout the column holds (Towler & Sinnott, 2013). The height of packing <math>Z</math> is given by the following equation:

<math>Z=\frac{L_m}{K_G*a*P}*\int\limits_{y_2}^{y_1} \frac{dy}{y-y_e}\,</math>

where <math>P</math> is the total pressure, <math>a</math> is the interfacial surface area per unit volume, <math>y_1</math> and <math>y_2</math> are the mol fractions of the solute in the gas stream at the bottom and top of the column respectively, <math>G_m</math> is the molar gas flow rate per unit cross-sectional area, and <math>y_e</math> is the mole fraction of solute in the gas that would be in equilibrium with the liquid concentration.

The first half of the equation before the integral can be called the height of an overall gas-phase transfer unit <math>H_G</math> and the second part of the equation is the number of overall gas-phase transfer units or <math>N_G</math>. Using these definitions the above equation can be simplified to

<math>Z=H_G*N_G</math>

These equations assist in sizing an absorption column (Towler & Sinnott, 2013).

==== Film Absorber ====
The final absorber the film absorber is generally used in the case where the heat of absorption must be removed. The film absorber operates by sending the gas and solvent through a heat exchanger where the solvent creates a thin film on the walls of the tubes and the gas flows through the interior allowing for solute transfer. The good heat transfer present in a film absorber makes it preferable for situations where low temperatures are required for a high recovery of the solute (Schmidt 2012).

===Industrial Absorption Processes===
An industrial example is lean oil absorption, which is used to separate nitrogen and other impurities from natural gas. A lean oil is contacted with low quality natural gas, and the methane is selectively absorbed by the lean oil, leaving the impurities behind. The methane is subsequently regenerated from the rich oil as high quality natural gas (Petrogas Systems, 2014).

Other common industrial practices of absorption come from sour gas treatment. Amine gas treating is used to remove hydrogen sulfide or carbon dioxide from gas streams via a reversible chemical absorption. In amine gas treating the sour gas is fed to the bottom an absorber where amine solution is fed to the top along with any necessary make up water. The sour gas components are absorbed into the amine via a chemical absorption method. Sweet gas leaves the top of the absorber whereas the amine out of the bottom, now rich with acidic components is sent to a regenerator where the acid gas components are stripped and the acid gas is generally sent to a flare whereas the amine now lean again is recycled back into the first absorber (Miller & Zawacki, 1978). Figure 5 below shows the typical setup of an amine plant. Another type of sour gas treatment that uses absorption is Merichems LO-CAT process which uses a chelated iron to remove hydrogen sulfide from feed gas in the absorption column (Merichem 2015).
[[File:AmineTreating.png|thumb|400px|center|Figure 5. Amine Gas Treating Plant Schematic]]

==Stripping==
This process separates solutes from solvents (often after absorption, to purify the solvent so that it can be recycled to an absorber). Stripping will depend on the vapor and liquid flow rates, as well as the temperature and pressure of the column. There is a temperature drop down the column, so columns generally have either an increased operating temperature or decreased operating pressure.

The stripping factor of component i is

<math>S_i=K_iV/L</math>

where <math>K_i</math> is the vapor/liquid equilibrium ratio, <math>V</math> is the vapor flow rate entering the column, and <math>L</math> is the liquid flow rate entering the column, will determine how much of solute i will be stripped from the liquid into the vapor phase (Peters & Timmerhaus, 2003). The usual range for the stripping factor is between 1.2 and 2.0, with a stripping factor of 1.4 being most economic.

An example of stripping in industry is the deodorization of food items such as oils. The oil is heated and allowed to trickle down the column while steam flows up from the bottom of the column. At the vapor-liquid interface, volatile components of the oil transfer to the steam and are carried off the top of the column, leaving a purified oil product (Alfa Laval, 2014).

==Bioseparations==
===Importance===
As our ability to manipulate and engineer biological systems improves, biological products are becoming an increasingly important source of therapeutics and fuels. The production of fuels from biomass via either the enzymatic breakdown of a feedstock or the secretion of usable lipids from algae is a promising new energy source. Additionally, enzymes, antibodies and other therapeutic proteins have been applied to the treatment of a wide range of diseases. Although each process requires its own set of separations, all follow the same basic format: separation of biomass, product isolation, and product purification (Belter et al., 1998). This section will provide examples of unit operations in each step. Ultimately, the choice of separation process and unit operations will depend on the specific process and product. The descriptions below are examples of the most common bioseparation operations within the general platform (Harrison et al., 2003).

Bioprocesses begin with fermentations or growth operations. In biofuel production processes, this may involve growing algae or breaking down corn or cellulosic biomass. For the production of therapeutics, mammalian or bacterial cells may be grown in a fermentor and the product secreted into the supernatant or harvested from the cells.

===Biomass Separations===
After fermentation and product production, the solid biomass must first be separated from the desired product. If the product is secreted from the cells, this can be done immediately after fermentation ends. If the product is not secreted, the cells must first be lysed.
Cell lysis is the process of lysing, or breaking, the cell in open. Mechanical lysis is the simplest, and involves physically breaking the cell either by mashing (think mortar and pestle) or blending the cells into a homogenous solution in a homogenizer. Chemical lysis is another method, achieved by introducing an osmotic shock or chemically degrading the cell membrane. Additional separation can be achieved by flocculation, which is the process of aggregating biomaterial by charge neutralization or bridging. These larger complexes are easier to separate from smaller molecules (Harrison et al., 2003).

The next step is removing the unwanted biomass from the product in solution. Separation by centrifugation or sedimentation are the most common, although filtration is sometimes also used for processes where a biomass cake is desired. Both methods utilize density differences to separate the product from the solid biomass (Towler and Sinnott, 2013).

====Sedimentation====

Sedimentation relies purely on the force of gravity, while centrifugation speeds the settling process by subjecting the cells to a centrifugal force. Sedimentation in a settling tank is the simplest method of solid-liquid bioseparation. In this process, biomass in a tank is simply allowed to settle to the bottom over time. While this process is inexpensive, requires little energy and can separate out large volumes of biomass, it generally requires long time periods and is only mostly in very large-scale processes where active centrifugation is difficult (Belter et al., 1998).

====Centrifugation====
Centrifuges are widely utilized across many processes, and thus a wide variety of scales and designs have been developed. Disk-stack centrifuges, in which the solid phase is deposited onto “shelves” in the center of the spinner and liquid phase is pushed to the outside, are some of the most commonly used centrifuges in industry. They are especially suited to biomass separation processes because they can be built on a large scale and are ideal for separating fine solids from liquids. [[File: Disk_stack_centrifuge_towler.png|frame|center|Fig. 6: Diagram of a disk-stack centrifuge (Tolwer et al, 1997).]] Tubular bowl centrifuges are also common and can reach separation efficiencies of up to 90%. Heavier products accumulate along the sides of the bowl, while the light phase flows out the top. They separate products by can be used both to separate solids from liquids and immiscible liquids, such as and oil product and an aqueous broth (Tolwer and Sinnott, 2013). [[File: tubular bowl centrifuge towler.png|frame|center|Fig. 7: Diagram of a tubular bowl centrifuge centrifuge (Tolwer and Sinnott, 2013).]]

Centrifugation scale-up is made easier by sigma analysis, which allows for the estimation of appropriate feed rates for different size centrifuges. The sigma factor is dependent on the inner and outer radius of the centrifuge, the angular velocity, and the sedimentation velocity of the solid particles being separated. It can be thought of as the characteristic cross-sectional area with units of [length]2. The sedimentation velocity can be calculated by

<math>v_g={\frac{2a^2(\rho-\rho_0)}{9\mu}}g</math>

where <math>v_g</math> is the sedimentation velocity, <math>a</math> is the cell or biomass particle diameter, <math>\rho</math> is the particle density, <math>\rho_0</math> is the fluid density, and <math>\mu</math> is the fluid viscosity. The volumetric flow <math>Q</math> can be estimated by

<math>Q=(v_g)(\Sigma)</math>.

The equality

<math>{\frac{\Sigma_1}{\Sigma_2}}={\frac{Q_1}{Q_2}}</math>

can be an easy way to estimate equivalent flow rates between a small-scale centrifuge 1 and larger centrifuge 2 (Harrison et al., 2003).

====Example: Centrifugation Scale-up====

You are trying to separate a cell of radius 0.4 <math>\mu</math>m with a density of 1.05 g/cm3 from broth of mostly water (density of 1 g/cm3 and viscosity of 0.01 g/cm s). The sigma factor of the centrifuge you are using is 1 x 106 cm2. A] What volumetric flow rate should you use? B] If you want to scale up the process to a centrifuge with <math>\Sigma</math> = 3 x 106 cm2, what flow rate would you use in the larger centrifuge?

Solution:
A] Using the equation for <math>v_g</math>, and being mindful of units, the sedimentation velocity equals 1.74 x 106 cm/s. The flow rate, then, equals

<math>Q=(1.74 x 10^-6)(1,000,000) = 1.74 cm^3/s = 0.104 L/min</math>.

B] Keeping in mind that for the same process, <math>v_g1 = v_g2,</math> and rearranging the sigma factor equality, the new flow rate is

<math>Q_2 = {\frac{\Sigma_2 x Q_1}{\Sigma_1}} = {\frac{(3 x 10^6)(0.104)}{1 x 10^6}} = 0.313 L/min </math>

===Product Isolation===
Liquid-liquid separation, to extract the product from the aqueous phase, is much less straightforward than liquid-solid extraction. Many methods - especially adsorption, filtration, and precipitation - are similar in principle to operations found in other, non-biological separations. The exact separations used depend on the nature of the product and the scale of the process. These processes are nearly identical to their non-biological counterparts, and their description is left to other sections.

Particular care needs to be taken with protein products because of their instability, and the selection of an appropriate solvent or adsorbent is crucial to a successful process (Harrison et al., 2003).

===Product Purification===
The final steps of protein purification and polishing remove any remaining contaminants and bring the concentration of product to an appropriate value for applications. Purification processes for food-grade and medical products can be extensive, as sterility and high purity are essential. Purification in fuel-producing processes may be less extensive, depending on the process. Chromatography and crystallization are two common steps in purification and are especially used in industrial scale protein production. Several different types of chromatography exist with the ability to carry out different types of separations.
Chromatography is similar to adsorption in that it relies on differences in affinity between solutes and a solid surface. A solution is eluted through a column containing a solid resin with various affinities for the substances in solution. In adsorption, the solutes are evenly saturated throughout the column. Chromatography differs in that solutes are deposited a resin phase before the column is flushed with an elution solvent specific that results in solutes eluted in bands.

==== Ion Exchange Chromatography ====

There are two main types of ion exchange columns—anion and cation. Anion exchange resins have a positive charge and are used to retain products with a negative charge. Cation exchange resins have a negative charge and are used to retain products with a positive charge. The pH of the elution buffer is change to force a specific solute to wash out, depending on whether the pH of the buffer is above or below the isoelectric point of the solute (Belter et al., 1998). This is especially useful for the separation of protein product (including antibodies), nucleic acids, and other charged molecules. When the solutes have sufficiently different isoelectric points, the pH of the buffer is manipulated to affect the solute charge and force the product to elute while the solute remains preferentially bound to the resin, or vice versa (Harrison et al., 2003). In general, the most strongly charged molecules will remain in the column for a longer period of time. Elution washes through the weakly bound ions before the more strongly bound ions. Different speeds of elution can be visualized as in figure 8.

[[File:chromatography.png|frame|center|Fig. 8: Illustration of product bands in an elution chromatography column (Belter et al., 1998).]]

==== Size Exclusion Chromatography ====

In gel filtration chromatography, small molecules are "trapped' by the porous resin and take longer to flow through the column. Larger products will elute first because the smaller molecules are better able to penetrate the resin. This forces them to take a much longer path through the column, which means it takes longer for them to elute. This operation is often used when there is a distinct difference in size between the desired product and other solutes.

==== Affinity Separations ====

Affinity chromatography is very similar to ion exchange chromatography in that the interactions between the material in the column and the molecules in the feed. The main difference is that affinity chromatography can rely on a great variety of types of interactions. Two very common types of affinity are exploited in affinity chromatography columns. The first is immunoaffinity. Proteins are specifically bound by antibodies which can be incorporated onto beads and used in chromatography. Antibodies are designed to bind only a single protein, so these interactions are considered to be highly specific. The protein can be eluted using a buffer that changes the pH or salinity in the column, which adversely affects binding (Hage, 1999).

Proteins can be separated from very complex and unrefined mixtures using this method because of the great specificity. Antibody-protein interactions are so specific that individual proteins can be isolated from blood samples as shown in the figure below. The main drawback of this method is simply that specific interactions between proteins and binding targets do not always exist. Antibodies exist for many proteins, but for others they must be customized or created using some sort of evolution process. This is not a trivial task and can require a significant capital investment (Hage, 1999).

[[File:Affinity_Chromatography_Example.PNG|frame|center|Fig. 9: Affinity purification separating fibrinogen from human plasma using an anti-fibrinogen antibody (Hage, 1999).]]

The other main type of affinity chromatography is based on protein specific tags and the molecules or surfaces to which they bind. One of the most common types of protein tags used is the polyhistidine tag. This tag consists of 6-8 consecutive histidine residues which can be added to the exterior of the desired protein product. The addition of this tag requires alterations to the coding sequence of the protein. The polyhistidine tag binds strongly with nickel and cobalt ions. The product with the tag can then be eluted with imidazole—a small molecule with the same structure as the functional group of the amino acid histidine. Imidazole will bind the cobalt and nickel ions more strongly than the histidine in the tag. Along with chromatography, protein tag interactions can be leveraged with the use of beads that can be deposited directly into the solution containing the protein of interest (Lichty et al., 2005).

Different proteins can be purified in this manner with varying levels of efficiency. There is a very high dependence on the size of the protein, electronic properties, and steric considerations. The figure below demonstrates varying separation efficiencies with proteins of different sizes (HisPur Cobalt Resin - Thermo Fisher Scientific, 2016).

[[File:HisTagPurification.PNG|frame|center|Fig. 10: Protein gel demonstrating the separation of proteins from a sample. The far left column represents the most efficient separating conditions using cobalt and the far right column represents a negative control with no purification performed (HisPur Cobalt Resin - Thermo Fisher Scientific, 2016).]]

Several other types of tag-bead interactions can be utilized in separations processes. Maltose Binding Protein is a small protein that can be added to a protein of interest. It binds strongly with beads coated in immobilized maltose and can be released by flushing with maltose. As MBP is a full sized protein that typically must be removed from the protein of interest in order for it to be used. In this case, the site specific TEV protease is often used cleave MBP from the protein of interest. In addition, under specific circumstances, other unique tags can be used and provide varying levels of specificity in separations. The Flag tag, 3x Flag tag, Glut tag, and Strep tag. While these are all commonly used, the polyhistidine tag is the most popular because it gives the highest level of specificity. Some of the main purification tags are compared in the table below (Lichty et al., 2005).

{| class="wikitable"
|-
! Tag
! Size (aa)
! Resin
! Eluting Agent
! Capacity (mg/ml)
! Cost/10 mg
|-
| His
| 6
| Talon
| Imidazole
| 5-14
| $18
|-
| Maltose Binding Protein
| 396
| Amylose
| Maltose
| 3
| $12
|-
| Glutathione S-Transferase
| 218
| GSH-Sepharose
| Glutathione
| 10
| $36
|-
| Strep II
| 8
| Strep-Tactin Sepharose
| Desthiobiotin
| 0.1
| $293
|-
| FLAG
| 8
| Anti-FLAG M2 MAb Agarose
| Flag peptide
| 0.6
| $1045
|}

Unfortunately, based on the contents of the solution, the characteristics of the protein, and the specific product being used, there are extremely large amounts of variation in yield and purity for each type of tag. Therefore, it is crucial to allocate time towards preliminary tests and scale-up experiments with the exact solution from which the protein of interest is being separated. The his tag is typically the most used because, as seen in the table above, it has one of the lowest costs and highest yields. In addition, because it is only a 6 amino acids, it often is not cleaved from the protein of interest which minimizes the number of steps required in the overall purification process (Lichty et al., 2005).

==== Crystallization ====

Crystallization, or the formation of solute crystals from a solution, is especially useful in biomolecule separations because it is possible to obtain a 99.9%+ product purity. In crystallization, a diluent is added to the homogeneous solution that reduces the solubility of the product to the point that it “falls out” of solution and crystallizes. It is similar to precipitation but results in the formation of crystals rather than unordered aggregates. Crystallization can be used on a laboratory scale for determining protein structure, on on the industrial scale for antibody and therapeutic protein productions. Batch crystallizers are often used in industry because of their simplicity and inexpensiveness compared to continuous crystallization (Harrison et al., 2003).

==Membrane Separation==
Membrane separation takes advantage of the selective permeability of membranes; they allow certain particles to pass through and selectively stop other, generally unwanted, particles. The component that passes through is called the permeate and the component stream that is rejected is called the retentate or concentrate. The applicability of membranes comes from the fact that their selectivity is determined by their pore size, which can be controlled during the creation of the membranes. Additionally, Membrane processes do not require heat meaning they generally require less energy than conventional separations technology such as distillation and crystallization. Membrane separations processes are generally classified as microfiltration, ultrafiltration, or nanofiltration depending on the size of the particles to be filtered out.

[[File:membrane techs.png|frame|Figure. Cutoffs for different membrane categories]]

===Membrane Selection, Construction, and Flow Geometries===
Membrane permeability and selectivity are the two most important factors to consider when selecting a membrane. For gas separations, the permeation of the gas is usually facilitated by the gas dissolving in the membrane on one side and then evaporating on the permeate side. Therefore permeability depend largely on the solubility of components in the membrane.

The two most commonly utilized membrane configurations are hollow fiber and spiral wound. Hollow fiber is generally the most commonly utilized module for gas separations. These are formed by gluing the two ends of the hollow fiber to a resin forming a closure. The fibers are housed in a shell much like a heat exchanger. The feed flows past thousands of tubes with the permeate flowing into the hollow tubes and out the closure. The retentate then flows out of the shell not having gotten through the membrane.

[[File:hollow membrane.jpg|frame|Figure. Hollow fiber membrane module]]

Spiral wound membranes are created by sealing two membrane sheets back to back on three edges to form a sort of pocket. This fourth open edge is then attached to a porous tube which allows permeate to go through it. Several membrane pockets are attached to a single tube and wrapped around in a spiral.

[[File:spiral membrane.jpg|frame|Figure. Spiral wound membrane module]]

Flow geometry is usually either dead-end geometry or cross flow geometry. In dead end, the fluid flow is normal to the membrane surface while cross flow is parallel to the membrane surface. Dead end geometry is usually used with hollow fiber membranes while cross flow is used with spiral wound membranes. Each geometry has advantages and disadvantages. Dead end geometry is generally cheaper to set up and therefore has lower initial capital costs. However, it is very vulnerable to membrane fouling, which reduces the effectiveness of the membrane. This is usually the geometry set up for small scale lab experiments. The tangential flow devices are more cost and labor-intensive, but they are less susceptible to fouling due to the sweeping effects and high shear rates of the passing flow. Most commercial industrial membrane separations are done using spiral wound cross flow membrane modules.

===Applications===
====Food Industry====
Due to the fact that MD can be conducted at relatively low feed temperatures, it was successfully tested in many areas where high temperature applications lead to degradation of the process fluids especially in food processing. It was demonstrated that MD can be used for the concentration of milk, for the recovery of volatile aroma compounds from black currant juice, and for the concentration of many other types of juices including orange juice, mandarin juice, apple juice, sugarcane juice, etc.
====Reverse Osmosis====
Reverse osmosis is the most widely used membrane separation process. In this process, fresh water passes through the membrane while dissolved salts and other solids are rejected and stay in the concentrate. In this process, feed water is pressurized in order to overcome the osmotic potential difference between the salty retentate and the fresh water desired. These processes are generally run using spiral wound membrane cylinders using a cross flow setup.

===Membrane Model===
The two most important components when considering different membranes are the permeability, which will determine flux through the membrane, and selectivity, which will determine what passes through the membrane and how much. The flux through a membrane is defined as:
<math> M_i = \frac{P_i}{δ}(p_{i,f} - p_{i,p})</math>
Where Mi is the molar flux of component i, Pi is the permeability of the membrane for component i, δ is the membrane thickness, and pi,f and pi,p are the partial pressures of component i on the feed side and permeate side respectively.
The average flux across a long cylindrical membrane such as the spiral wound module is given by:
<math> \int_0^Lm \frac{M_i,dx}{L_m}</math>
Where Lm is the length of the cylinder and x is length in meters

Membrane selectivity of the ideal separation factor is given as the ratio of the permeability of one substance over another as shown:

<math> S_(i,j) = P_i/P_j </math>

where Sij is the selectivity of the membrane for component i over j.

==Cyclones==
Centrifugal Separators, more commonly know as Cyclones, are one of the most widely used gas-solid separators. They are typically used for particles that are 5μm or larger, but if agglomeration occurs they can sometimes separate particles as small as particles as small as .5μm. These cyclones can be designed for high efficiency separations or for high throughput (Towler, 2012).

[[File:cyclone.png|thumb|center|Reverse-flow Cyclone (Towler, 2008)|500x300px]]

===Design===
Design Procedure
There are multiple design methods for cyclones. The deign method discussed below is the Stairmand method. This involved creating two designs one of which was for high throughput and the other which was designed for high efficiency. Using these base designs future cyclones can simply be estimated using scaling factors (Towler, 2012).

1. Select whether efficiency or performance is more important for your design

2. Determine the particle size distribution in your stream

3. Estimate the number of cyclones needed

3. Calculate the cyclone diameter

4. Calculate the scale up factor

5. Calculate the cyclone performance and overall efficiency

6. Calculate the cyclone pressure drop

7. Cost the system

====Particle Size====
The diameter of particles separated by cyclones are governed by this design equation. This equation is dependent upon the standards of the cyclone and how much you are trying to deviate from its standard operating conditions. Looking into the equation it shows that a larger internal cyclone diameter, higher flowrate, higher difference in density, and lower viscosity result in smaller particles being seperated. The whole term after d1 is known as a scale up factor. When designing a cyclone you would determine the flow rate, viscosity, particle size, and density difference would be known prior to using this design equation, so this would be used to optimize the diameter of the cyclone you were designing (Towler, 2012).

[[File:Towler 10.8.PNG|center|Reverse-flow Cyclone (Towler, 2012)|500x300px]]

d1=mean diameter of particle separated at the standard conditions, at the chosen
separating efficiency

d2= mean diameter of the particle separated in the proposed design, at the same
separating efficiency

Dc1 =diameter of the standard cyclone

Dc2 =diameter of proposed cyclone

Q1 =standard flow rate

Q2 =proposed flow rate

Δρ1 =solid-fluid density difference in standard conditions

Δρ2 =density difference, proposed design

μ 1=test fluid viscosity

μ 2=viscosity, proposed fluid.

[[File:Towler10.44.PNG||thumb|center|(a) Standard Cyclone Dimensions (b) High Gas Rate Cyclone (Towler, 2012)|500x300px]]

====Efficiency====
Efficiency is the percentage of total solids that are removed from the stream The graphs below were experimentally for the two standard designs Stairmand Method. This graph has the efficiency vs the particle size.

[[File:Towler10.45.PNG||thumb|center|Performance curves, standard conditions. (a) High-efficiency cyclone performance curves, standard conditions. (b) High gas-rate cyclone (Towler, 2008)|600x400px]]

If you have to scale the efficiency it can be estimated by the point where a horizontal line from the efficiency of the original particle size intersects the new particle size. A demonstration is on the graph below.

[[File:Towler 10.46.PNG||thumb|center|Scaled Performance Curve(Towler, 2008)|500x300px]]

====Pressure Drop====
One of the most important design factors in a cyclone is its pressure drop. The pressure drop occurs because entry and exit losses, kinetic energy losses, and friction losses that occur inside the cyclone. The pressure drop can be calculated using the following equation.

[[File:Towler 10.9.PNG|center|Reverse-flow Cyclone (Towler, 2008)|500x300px]]

The variables are defined as follows:
DP = cyclone pressure drop, millibars;
rf = gas density, kg/m3
;
u1 = inlet duct velocity

u2 = exit duct velocity

rt = radius of circle to which the center line of the inlet is tangential

re = radius of exit pipe

φ =factor from figure below

ψ = fc(As/A1)

fc = friction factor

As = surface area of cyclone exposed to the spinning fluid

A1 = area of inlet duct

[[File:Towler_10.47.PNG|thumb|center|Cyclone Pressure Drop Factor (Towler, 2008)|600x400px]]

====Cost Estimation====
The following values are represented in 2002 dollars. Cost is mostly a function of the amount of gas that need to be passed through the cyclone and the units are in standard meters cubed per second (sm3/sec).

The costs can be broken down into three categories:

Capital Costs: $4600-$7600 sm3/sec

Operations and Maintenance: $1500-$18,000 sm3/sec

Annualized Cost: $2800-$29000 sm3/sec

This all comes to $.47-$440 per metric ton of pollutant removed. The high variation in these costs have to do with smaller units being much less efficient per unit of pollutant removed. Furthermore the cost is also effected by the amount of particulates that are in the gas to begin with. For example a gas with 10 g/sm3 and air with 100sm g/sm3 would have similar pumping costs, but very different cost per unit of solids removed.

==Other Separation Processes==
===Extraction===
Liquid-liquid extraction is a process for components with overlapping boiling points and azeotropes. The process requires a solvent such that some of the components of the mixture are soluble, and then the components will be separated based on this solubility in the liquid. This process can operate at moderate temperatures and pressures, so is not very energy intensive. However, a distillation column is required to extract the solvent for recycle. More recently, supercritical fluids have replaced liquid solvents in some processes for L/L extraction, due to the solute’s ability to more rapidly diffuse through them. The issue with these fluids, however, is that they must be operated at extremely high pressures and temperatures, increasing both capital and operating expenses of the process (Peters & Timmerhaus, 2003).

===Crystallization===
This process recovers solutes that have been dissolved in solution. The resulting product is in the solid phase. Depending on the material properties of the solute and solvent, the solute is recovered by precipitation after cooling, removal of solvent, or adding precipitating agents. Crystallizers are designed based on phase equilibria, solubilities, rates and amounts of nuclei generated, and rates of crystal growth. Every crystallization process is a unique system, so plant evaluation is usually required before complete implementation. Crystallization can be performed in both batch and continuous processes, and design features can control crystal size to an extent (Peters & Timmerhaus, 2003).

===Membrane Separation===
This separations process uses selectively permeable membranes to separate components in a mixture. Typically, one of the components will freely pass through the barrier while the other components will not. The stream that passes through the membrane is the permeate and the stream that does not pass is the retentate. The driving force behind this separation is a pressure gradient. Membrane separation is beneficial because it can separate mixtures at the molecular and small particle level. Furthermore, there is no phase change required so the energy input is low. Limitations of this process include achieving high product purity, incompatibility with certain stream components, low operating temperature, and low flow rates. Although membrane separation is generally not scaled up, examples of scaled-up membrane separation include seawater desalination and hydrogen recovery (Peters & Timmerhaus, 2003).

===Adsorption===
Adsorption involves an adsorbent and adsorbate. The adsorbent is typically a solid, and will typically separate the adsorbate from the stream. This process usually includes a desorption step that regenerates the adsorbent for further use. Raising the temperature or increasing the concentration of the adsorbate can reverse the adsorption process. Although the recycle of the adsorbent is a very economic design feature, the downside of this step is that it results in a cyclic process, which introduces complexity to the overall process. Industrial applications of this process are for bulk separations and gas purification. The adsorption/desorption process in these situations involves a large amount of heat transfer, which design engineers must take into account when sizing and selecting equipment material (Peters & Timmerhaus, 2003).

===External Field/Gradient Separation===
These separations use external force fields or temperature gradients to separate responsive molecules or ions. The use of these processes is fairly limited to a few specialized industrial applications (Peters & Timmerhaus, 2003).

===Settling and Sedimentation===
In settling processes, solid particles or liquid drops are separated from a stream by gravity. The stream can be in either the liquid or gas phase. For vapor-liquid mixtures, flash drums are generally used to separate the mixture. The velocity of the vapor must be less than the settling velocity of the liquid drops for this separation to occur. For liquid-liquid separation, the horizontal velocity of the fluid must be low enough to allow the low-density droplets to rise to the interface and the high-density droplets to move away from the interface and coalesce. In sedimentation, the result of the process is a more concentrated slurry. Typically a flocculating agent is used to aid in the settling process. One way to perform this separation is to use a cone-shaped tank with a slowly revolving rake that scrapes and moves the thickened slurry to the center of the cone for removal (Peters & Timmerhaus, 2003).

[[File:Example.jpg]]

====Clarifiers====
[[File:Circular_Clarifier.png|300px|thumb|bottom|Figure 9: Circular clarifier with some components labelled.]] [[File:Rectangular_Clarifier.png|300px|thumb|bottom|Figure 10: Rectangular clarifier with some components labelled.]]

Clarifiers are one of the methods used for the continuous removal of particulate solids from liquids through sedimentation by gravity. Applications include process water pretreatment, waste water treatment, and drinking water purification. Historically, clarifiers were originally developed to limit nutrient input into surface water due to fear of eutrophication. Today, they have a number of uses, particularly in wastewater treatment processes, metal removal, disinfection, and membrane pretreatment. The process helps removed dissolved solids, silt, and undesirable metals from the water, making it more suitable for downstream processes as well as human consumption (Wilson, 2005).

Clarifiers are typically used in conjunction with coagulation or flocculation agents, which promote dissolved particles to join into clumps and settle out of solution (Towler and Sinnot, 2012). Clarifiers typically consist of a large circular tank with a rotating rake at the base which scrapes settled solids towards the center. In the case of a rectangular clarifier, they are scraped to one side. Diagrams of both are represented in figures 9 and 10, respectively (NMED Surface Water Quality Bureau, 2015). Separated solids are allowed to settle to the bottom of the tank as a sludge, whereupon they are collected by the rake and disposed of properly. In the case of floating contaminants, it is possible for the clarifier to include a skimmer as well.

Clarifier efficiency varies with certain factors, including the settling characteristics of solids removed and the surface overflow rate of the tank. Clarifier efficiency can be found using the following relation:

<math>\begin{align}
E_{TSS} &= E_{TSSmax}\left ( 1 - e^\frac{\lambda}{SOR} \right )
\end{align}</math>

where <math>E_{TSS}</math> is the efficiency of total suspended solids (TSS) removal, <math>E_{TSSmax}</math> is the maximum possible efficiency, <math>\lambda \left [\frac{m}{d} \right ]</math> is the settling constant, and <math>SOR \left [\frac{m^3}{m^2 d} \right ]</math> is the surface overflow rate. The effect of flocculation chemicals on TSS can be seen in figure 11. However, it should be noted that chemical addition will increase sludge quantity and may have an adverse effect on plant aesthetics, which increases maintenance costs (Wilson, 2005).

[[File:Chem_Addition.png|200px|thumb|bottom|Figure 11: The effect of flocculating agents on total suspended solids removal in clarifiers.]]

=====Lamella Clarifiers=====

Lamella clarifiers use inclined plates in order to maximize the settling area for solids. Solids continue to settle into a hopper at the bottom of the tank while clarified water exits up through the inclined plates. This allows for the design of a smaller tank, which leads to large savings in capital costs. A lamella clarifier is pictured in figure 12.

[[File:Lamella_Clarifier.png|300px|thumb|bottom|Figure 12: A lamella clarifier with components labeled.]]

Typically, inclined plates are installed at an angle of 45 to 60 degrees and spaced 40 to 120 mm apart, which increases effective settling surface area by a factor of 6 to 12 compared to traditional clarifiers. For effective use, it is recommended that the Reynolds number be below 2000, Froude number higher than 10-5,and detention time be longer than 3 to 5 minutes. For this implementation, the equations are as follows:

<math>\begin{align}
N_{Re} &= \frac{VR}{\nu} \\
N_{Fr} &= \frac{V^2}{Rg}
\end{align}</math>

where <math>R</math> refers to the hydraulic radius, which is the cross-sectional area of the lamella, <math>V</math> is the liquid velocity, <math>\nu</math> is the kinematic viscosity, and <math>g</math> is the gravitational constant (Wilson, 2005).

=====Advantages=====

Clarifiers offer a proven, relatively inexpensive solution for solids removal. The chemical coagulants used are cheap and provide a low operating cost as well as simple maintenance. Construction is typically simple, leading to low capital costs and equipment that is easy to accommodate and maintain. Their design is also flexible, with various options such as skimmers and scrapers offering increased removal efficiency (Wilson, 2005). Operation of clarifier tanks also has lower energy requirements than membrane filtration for solids removal, given that most of the separation is aided by gravity. Water exiting clarifier units has a silt density index (SDI) averaging 4.0, which is low enough for further membrane treatment such as reverse osmosis (Prihasto, 2009).

=====Disadvantages=====

Clarifiers necessitate low turbulence to prevent resuspension of solids. This essentially requires a low entrance velocity, which can limit the production rate of certain processes or call for more clarifier units, which would drive up costs. Furthermore, clarifiers require frequent cleaning before sludge becomes too difficult to remove and reduces effectiveness. In the case of lamella clarifiers, sludge buildup on the inclined plates results in uneven flow distribution which could harm efficiency (US EPA, 2003). For this reason, maintenance requirements for lamella clarifiers are higher, but they can be reduced through the implementation of removable plates (Wilson, 2005). Clarifiers also only remove solids, so pH will not be affected, leading to the need for further pH adjustment (NMED Surface Water Quality Bureau, 2015).

=====Clarifier Design Calculations and Typical Design Values=====

======Detention Time======

Detention time (DT) is the time is takes for a unit of water to travel from the inlet of the clarifier unit to the outlet. During typical operations, the design value for this is 2 to 3 hours.

<math>\begin{align}
DT &= \frac{Tank\ Volume}{Influent\ Rate}
\end{align}</math>

======Surface Overflow Rate======

Surface overflow rate (SOR) measures the flow into the clarifier per square foot of surface area. Typical design values are 400 to 800 gal/day/sq. ft.

<math>\begin{align}
SOR &= \frac{Volumetric\ Flow\ Rate}{Surface\ Area}
\end{align}</math>

======Weir Overflow Rate======

Weir overflow rate (WOR) describes the flow in gallons per day per linear foot of weir. Typical values are 10,000 gal/day/ft.

<math>\begin{align}
WOR &= \frac{Volumetric\ Flow\ Rate}{Weir\ Length}
\end{align}</math>

======Solids Loading Rate======

Solids loading rate (SLR) describes the mass of solids in the clarifier influent per square foot of surface area. This value should not exceed 30 lbs/day/sq. ft.

<math>\begin{align}
SLR &= \frac{Solids\ Mass\ Flow\ Rate}{Surface\ Area}
\end{align}</math>

===Flotation===
Flotation is a process designed for specific solid-solid mixtures. It works by generating gas bubbles in a liquid that attach to selected solid particle. Afterwards, the particles rise to the liquid surface where they are removed by an overflow weir or mechanical scraper. The separation depends on the surface properties of the particles and its preference to attach to the gas bubbles. To meet the necessary requirements of the flotation process, a number of additives can be used to control things like the pH of the liquid-solid mixture, the activity of the solid surface, and the froth that can assist in separation. The bubbles can be produced by gaseous dispersion, dissolution, or electrolysis of the liquid (Peters & Timmerhaus, 2003).

===Centrifugation===
This process is similar to external field separation in that an external force field is applied to separate a mixture. When gravity separation is too slow due to particle densities, particle size, settling velocity, or the formation of an emulsion, centrifugation is commonly used. Centrifugal force increases the total force acting on the particle and results in faster separation times. This process is generally used to separate solids from liquids, however it can also be used to separate two liquids with very different densities (Peters & Timmerhaus, 2003).

===Drying===
Drying is performed to remove liquid from a liquid-solid mixture and produce a dry solid. Water is most often the liquid removed, but organic liquids are removed from solids on occasion as well. The heat required to vaporize the liquid is usually obtained by a series of gas-solid contacting devices. Feed condition and temperature sensitivity of the solid dictate the type of contacting device that is used. There are two groups of dryers that differ by the dependence of either mechanical means or fluid motion for gas solid contact. Another feature of dryers is to use either direct (hot gas) or indirect (conductive surface) heating (Peters & Timmerhaus, 2003).

===Evaporation===
Evaporators separate solvents from a solution by evaporation. The difference between evaporation and distillation is that evaporation requires the solute be nonvolatile. Because of this, a high separation can be achieved with one stage. Evaporators are essentially reboilers, so evaporation is a very energy-intensive process with a high thermal economy (Peters & Timmerhaus, 2003).

===Filtration===
Filtration is a process that separates a mixture of solid in a liquid or gas by passing the mixture through a porous medium in which the particles do not pass. Filtration is done by either cake filtration (particles found on the surface of the filter) or depth filtration (particles found within the filter). Cake filtration is generally performed with a cloth as the filtration medium (Peters & Timmerhaus, 2003).

==Conclusion==
Separation is a key part of most chemical processes, and there is a great variety of techniques to perform separation of compounds based on size, volatility, charge, and many other features. A common technique with which the process engineer should be familiar is distillation, but he or she should also be aware of the other available options. Some techniques may be less expensive, less energy-intensive, or more effective than distillation, depending on the specific separation problem. Therefore, the separation strategy should be carefully considered.

==References==
Belter PA, Cussler EL, Hu WS. Bioseparations: Downstream Processing for BIotechnology. New York: John Wiley; 1998.

Biegler LT, Grossmann IE, Westerberg AW. Systematic Methods of Chemical Process Design. Upper Saddle River: Prentice Hall; 1997.

Danckwerts P (1965) The Absorption of Gases in Liquids. Pure and Applied Chemistry UK 10:625-642.

Development Document for the Final Effluent Limitations Guidelines and Standards for the Metal Products and Machinery Point Source Category (Report). US Environmental Protection Agency. 2003.

Erwin, D. Industrial Chemical Process Design. New York: McGraw Hill, Professional Engineering; 2002.

Hage, David S. 1999Affinity Chromatography: A Review of Clinical Applications. Clinical Chemistry 45(5): 593–615.

Harrison RG, Todd P, Rudge SR, Petrides, DP. Bioseparations Science and Engineering. New York: Oxford University Press; 2003.

Hage, David S. 1999Affinity Chromatography: A Review of Clinical Applications. Clinical Chemistry 45(5): 593–615.

HisPur Cobalt Resin - Thermo Fisher Scientific N.d. https://www.thermofisher.com/order/catalog/product/89964, accessed February 21, 2016.

Lamella Plate Clarifier. Hydro International Web site. Available at: http://www.hydro-int.com/uk/products/lamella-plate-clarifier?s=0&r=uk. Accessed February 2, 2016.

Lean Oil Absorption. PetroGas Systems Web site. Available at: http://petrogassystems.com/technology/natural-gas-processing-and-dew-point-control/lean-oil-absorption. Accessed February 19, 2014.

Lichty, Jordan J., Joshua L. Malecki, Heather D. Agnew, Daniel J. Michelson-Horowitz, and Song Tan 2005Comparison of Affinity Tags for Protein Purification. Protein Expression and Purification 41(1): 98–105.

Merichem Gas Technologies. ®LO-CAT PROCESS available at http://www.merichem.com/images/casestudies/Desulfurization.pdf Accessed 6 Feb. 2015.

Miller L.N. & Zawacki T.S. , US 4080424, "Process for acid gas removal from gaseous mixtures", issued 21 Mar 1978, assigned to Institute of Gas Technology

NMED Surface Water Quality Bureau, New Mexico Water Systems Operator Certification Study Manual, New Mexico Environment Department, 2015.

Peters MS, Timmerhaus KD. Plant Design and Economics for Chemical Engineers. 5th ed. New York: McGraw Hill; 2003.

Prihasto, N; Lui, Q; Kim, S. Pre-treatment strategies for seawater desalination by reverse osmosis system. 2009; 249(1): 308-316. doi:10.1016/j.desal.2008.09.010

Schmidt Eberhard (2012) Waste Gases, Separation and Purification. Ullman’s Encyclopedia of Industrial Chemistry Germany 2:174-181.

Seider, W.D., Seader, J.D., and Lewin, D.R. (2004). ''Process Design Principles: Synthesis, Analysis, and Evaluation.'' New York: Wiley.

Stripping Column. Alfa Laval Web site. Available at: http://www.alfalaval.com/solution-finder/products/soft-column/Documents/Stripping%20Column.pdf. Accessed February 19, 2014.

Towler G, Sinnott R. Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013.

Turton, R.T., Bailie, R.C., Whiting, W.B., and Shaewitz, J.A. (2003). ''Analysis, Synthesis, and Design of Chemical Processes'' Upper Saddle River: Prentice-Hall.

Wankat, P.C. (2012). ''Separation Process Engineering.'' Upper Saddle River: Prentice-Hall.

Wilson, T.E., Clarifier Design, 2nd Ed., McGraw-Hill: New York, 2005.