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Natural Gas to Hydrogen (H)

2015-03-13T20:29:42Z

Mzuckerman:

Team H Final Report

Authors: Nick Dotzenrod, Samson Fong, Vince Kenny, Matthew Leung, Matthew Nathal, John Plaxco, Spencer Saldaña, Micah Zuckerman, Erik Zuehlke

Instructors: Fengqi You, David Wegerer,

March 13, 2015

__TOC__
==Executive Summary==

==Introduction==

As the need to become less reliant on fossil fuels grows, the idea of widespread hydrogen fuel cells comes closer to being a reality. Hydrogen-fueled vehicles produce no direct pollution, making them environmentally friendly. However, fossil fuels are often used in the processes for the hydrogen production and the energy content of the produced hydrogen is less than that of the fossil fuels used (Clay et al. 2004; Actual Worldwide, 2007). To successfully implement a change to hydrogen fuel, high capacity plants must be constructed to meet increased demand. With the increased production of natural gas and the corresponding drop in price, steam reforming is a profitable method to produce hydrogen. The goal of this project is to design, optimize, and economically analyze a steam reforming process that will produce 99.999% hydrogen from a feedstock of natural gas. Waste products such as carbon dioxide, waste water, and monoethanolamine will be treated to minimize environmental impact.

==Technical Approach Taken==

After investigating process alternatives for each of the design stages, the final technologies chosen to achieve the specified design criteria were chosen by comparing the advantages and disadvantages as well as considering factors unique to the proposed design. The major decision was to use steam reforming to generate the hydrogen. The specific type of steam reformation chosen was conventional steam reforming due to capacity constraints. A nickel catalyst was chosen due to its low cost. To treat the CO waste, both methanation and a water gas shift were selected to ensure product purity. Amine scrubbing was picked to separate carbon dioxide from the product stream. Finally, pressure swing adsorption was chosen to handle separation of the methane from hydrogen. Several process design options were investigated to determine the most economically favorable and environmentally neutral design. The three major processes used for hydrogen production evaluated were steam reforming, partial oxidation, and autothermal reforming. Various technology options including reactor types, catalysts, and separation methods were compared to find the optimal combination of reliability, sustainability, profitability, and versatility.

The Peng-Robinson fluid package was chosen for use in the HYSYS simulation because it is ideal for describing hydrocarbon systems such as the steam methane reforming plant, where most components are gases or are nonpolar. Furthermore, Peng-Robinson is reliable over a large range of operating temperatures and pressures, which encompass those within our system. The amine plant portion of the simulation uses the amine package. This package is based on the Kent-Eisenberg model specifically designed for modeling removal of CO2 and H2S by amines.
Difficulties encountered modeling adsorption processes in HYSYS lead to the use of component splitters to model the feed pretreatment and pressure swing adsorption (PSA) steps.

==Steam Reforming Process==

The pre-treatment zone desulfurizes the hydrocarbon feed. Then, it is combined with superheated process steam. Afterwards, this mixture is heated and passes through reformer tubes filled with catalyst (Hydrogen Production, 2004). The reformer tubes are also heated which causes the hydrocarbon and steam mixture to react, yielding carbon monoxide, carbon dioxide, and hydrogen (Steam Reforming, 2014). The water gas shift then converts the carbon monoxide to carbon dioxide and generates hydrogen. The plant is normally operated with a higher steam content than required to prevent elemental carbon from being deposited in the catalyst bed (Steam Reforming, 2014). The pressure swing adsorption will isolate the hydrogen desired from any residual gases that are later used as fuel. Figure 1 below shows the process flow diagram of the process. The stream and vessel number descriptions are found in Appendix A.

==Process Flow Diagram==

[[File:hydrogenflowsheet.png|thumb|800px|center| Figure 1. Process Flow Diagram of the Hydrogen Reforming Process]]

==Process Alternatives==

Three primary techniques are used to produce hydrogen from a hydrocarbon source, such as methane. The techniques are: steam reforming, autothermal reforming, and partial oxidation.
Each of these techniques has advantages and disadvantages which must be considered in choosing a final design.

:;Steam reforming
::Advantages: industrially mature, no oxygen needed, lowest process operation temperature

::Disadvantage: large volume of air emissions

:;Autothermal reforming
::Advantage: low methane slip

::Disadvantages: requires air/oxygen, little commercial experience

:;Partial oxidation
::Advantages: no catalyst needed, needs less desulfurization, low methane slip

::Disadvantages: high process temperatures, process has high a degree of complexity, poor hydrogen to carbon monoxide ratio

Steam reforming is the most appropriate technique for the proposed design due to its high degree of industrial maturity and safe operating conditions. Although partial oxidation and autothermal reforming each have distinct advantages, the lack of maturity for both of these processes would add a high degree of risk to the overall plant design. Furthermore, the lower operation temperature of steam reforming will result in a safer process (Holiday et al. 2007).

The five methods of steam reforming that would provide the necessary purity for hydrogen cars are the conventional box method, conventional can method, the compact steam methane reformers, plate-type steam methane reformers, and membrane reactor steam reforming.

:;Conventional box steam reforming
::Advantages: Conventional reforming can handle a large throughput such as the 100 MMscfd. The desired 99.99% purity is achieved after the water-gas shift and pressure swing adsorption (Ogden, 2002).6

::Disadvantages: Temperatures of 750-800 °C are needed in the reactor. The utility costs associated with this method are high (Schoerner et al. 2011).7

:;Conventional can steam reforming
::Advantages: Can be quickly constructed for fast start up.

::Disadvantages: Only cost efficient at throughputs under 10 MMscfd so multiple would need to be used (Schoerner et al. 2011).

:;Compact steam methane reformers
::Advantages: Concentric annular catalyst beds can produce the 99.999% purity which is more pure than required.

::Disadvantages: Produce 1 MMscfd and, therefore, this alternative is unable to handle the throughput of the plant being designed (Ogden, 2002).

:;Plate-type steam methane reformers
::Advantages: Plate-type steam reformers are less expensive, more efficient, and have faster startups than conventional methods.

::Disadvantages: The catalysts used are extremely sensitive to carbon deposition and sulfur poisoning. Implementing this process is more risky as it is not yet being used commercially (Ogden, 2002).

:;Membrane reactors for steam reforming
::Advantages: This method produces extremely pure product.

::Disadvantages: Membrane reactors require temperatures 200 °C greater than conventional methods and does not work effectively at the large scale desired (Ogden, 2002).

After taking all the processes into consideration, the conventional box steam reformer is the most appropriate choice for the plant being designed as it will yield the desired purity and can produce the 100 MMscfd desired.

==Technology Options==

===Reactor and Catalyst Alternatives===

Of the several process alternatives described above, steam reforming is the most developed technology for hydrogen production. However, there are different ways to carry out the process of steam reforming. The alternatives focus on the choice of catalyst and the configuration of the catalyst in the reactor.

Generally, catalysts can be categorized as non-precious metal (such as nickel) and precious metal (such as platinum and rhodium). The cost of non-precious metal catalysts is substantially lower, but the catalyst is also far less effective. Nevertheless, conventional steam reforming plants usually use nickel catalysts because heat and mass transfer effects generally dominate the reaction kinetics with an effectiveness factor as low as 5% (Adris and Pruden, 1996).

To reduce the heat and mass transfer effects, some processes attempt to reduce the particle size of catalysts and increase the reaction area by using microchannel-based reactors (Wang et al. 2004). As the size of the catalyst particle decreases, the apparent kinetics of reaction converge to the intrinsic kinetic of the catalyst. As a result, the more expensive precious metal catalysts are favored (Rostrup-Nielsen, 2003). Currently, researchers are developing cobalt-based catalysts to mitigate the cost of rhodium based catalysts (Song et al. 2007).

===Carbon Monoxide Removal Alternatives===

The steam reforming process produces approximately 5% of carbon monoxide in the product stream. Water-gas shift and methanation are two most common method to improve purity (Song, 2002). Water-gas shift reacts carbon monoxide and water to form hydrogen and carbon dioxide. In order to prevent the oxidation of hydrogen, precise control of air input is needed. Conventionally, the product stream is passed through a high temperature reactor because the reaction is substantially faster at higher temperature, but it is then passed through a low temperature reactor to improve hydrogen production (Hoogers, 2003). TeGrotenhuis et al. (2002) have demonstrated that a single reactor with a gradient temperature that spans both ends of the temperature extremes can be used in order to reduce the utility used. While copper catalysts are most prevalent, increased selectivity can be achieved with higher cost molybdenum carbide or platinum-based catalysts (Patt et al. 2000; Chandler et al. 200).

Another process to reduce carbon monoxide is methanation. Similar to water-gas shift, methanation reacts carbon monoxide and hydrogen to form methane and water. Methanation reactors are generally simpler without the need for a precise air stream. However, the reaction fundamentally consumes hydrogen at a very high rate (3 moles of hydrogen are consumed for 1 mole of carbon monoxide consumed) (Hoogers, 2003). A combination of both methods is generally used for high purity hydrogen production, which is appropriate for the proposed design. As such, both a water-gas shift as well as methanation will be utilized.

===Acid Gas Removal===

After the water gas shift, the product stream will be enriched in carbon dioxide and methane in order to remove carbon monoxide. The two stack gases will be separated in two steps. Carbon dioxide is an acid gas, which is generally removed by adsorption using an amine solvent (Kohl and Nielson, 1997). The product stream is passed through an absorber with the amine solvent where the carbon dioxide is stripped. The amine, now rich in carbon dioxide is then passed through a regenerator where the acid gas is separated from the amine (Miller and Zawacki, 1978). This process is referred to as amine scrubbing.

Other technologies that are used for acid gas removal include using a membrane to separate the acid gas and vacuum swing separation. Membrane gas separation however is still not mature enough for widespread usage today (Baker, 2002). Another alternative is to use vacuum swing adsorption. The feed gas is passed through a parallel adsorption bed where the carbon dioxide is adsorbed. The bed is then rinsed with a stream rich in carbon dioxide to remove any methane or hydrogen trapped in the bed. The pressure within the adsorber is then lowered to further the separation of the gases (Wang, 1989). Due to the complexity of the vacuum swing adsorption, and the immaturity of membrane separation, using an amine solvent for amine scrubbing is the method of choice for removing the carbon dioxide.

===Methane Removal===

Two potential methods of separating methane from hydrogen and water are membrane gas separation and pressure swing adsorption. Membrane gas separation is the less mature and reliable of these technologies. It involves the use of either a porous or non-porous membrane in order to selectively separate gas components primarily on size and affinity. Synthetic membranes are contained in either hollow fiber bundles, spiral wound bundles, or plates (Yun and Oyama, 1989).

Pressure swing adsorption (PSA) selectively adsorbs gas components by rapidly pressurizing and depressurizing a gas around adsorbents. Species are adsorbed at high pressure and the adsorbents can be regenerated by and series of purging and desorbing at low pressures (What is PSA, 2015). Common adsorbent materials are zeolites, molecular sieves, silica gel, alumina, and activated carbon. PSA operates at near-ambient temperatures and is the most widely accepted method of large scale hydrogen purification (Strasse, 2012).

==Key Processes and Products==

There are seven major processing steps involved in the steam methane reformation process as designed. These major processing stages are: the initial heating of feed, the steam reformation reactor, the high and low temperature water shift reactors, the amine plant, the methanation reactor, the gas compression and cooling train, and the pressure swing adsorption unit.

From the heating of the feed, the combustion of hydrocarbons to provide heat yields carbon dioxide as a major byproduct. The product of this process is the heated feed which contains methane and water. The main product of the steam reformation reactor contains methane, carbon dioxide, carbon monoxide, water, and methane. This product goes to the shift reactors where the outgoing product will contain less carbon monoxide and more carbon dioxide than that from the steam reformation reactor, but will otherwise contain similar components. The product from the shift reactors is sent to an amine plant where the acid gas (Carbon Dioxide) will be removed and vented to the atmosphere as waste product, resulting in a product stream containing carbon monoxide, methane, hydrogen, and water. This stream is sent to the methanation reactor where the carbon monoxide is broken down into oxygen and carbon which yields a product stream containing water, methane, and hydrogen. This stream is then sent to the compression system where the pressure is increased so that pressure swing adsorption may be used to separate the methane and hydrogen. From the pressure swing adsorption unit, the purified hydrogen is removed and sent to a storage tank to be sold later and the methane captured is then recycled back to the gas fired heater. For more details on the products in each flow, see Appendix A.

==Equipment Sizing==

Two sources were used to size the equipment and estimate capital costs. The first was the Aspen HYSYS integrated economics evaluator. Carbon steel was the chosen material for use in the furnace, pumps, and pipes. All components were sized with a 10% pressure safety margin or 25 psi greater than the operating conditions (whichever value was higher) in order to provide a potential buffer in case of an increase in temperature or pressure. This resulted in moderately higher prices, but no expense was spared due to safety concerns. The reactors were sized using methods described in Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design.

Since the PSA columns could not be modeled in HYSYS, individual suppliers were contacted to achieve the required plant capacity. Using empirical correlations, the capital cost was estimated for six PSA columns, shown in Appendix B. The equipment sizes (diameter, etc.) for certain equipment utilized in the process are shown in Appendix C.

==Economic Analysis==

Total Capital Cost: $6.26 x 107

Table 1 shows the annual operating costs for the process. More detail on how these costs were calculated can be found in Appendix D.

{| class="wikitable" style="margin: auto;"
!colspan="2"| Table 1: Annual Operating Costs
|-
! Process Expense
! Yearly Cost ($)
|-
| Waste Treatment
|1.86 x 106
|-
| Utilities
|2.98 x 107
|-
| Labor
|1.80 x 107
|-
| Maintenance
|1.88 x 106
|-
| Feedstock Purchase
|3.07 x 108
|}

Total Annual Operating Cost: $3.59 x 108, assuming the process will be operated for 24 hours a day, 360 days a year (to allow for holiday vacations).

The water will be sold at $0.00207 per kilogram (Freeport Water, 2014). Methane will be purchased for $0.226 per kilogram (U.S. EIA, 2015). MEA will be purchased for $0.50 per kilogram (Alibaba, 2015a). Hydrogen fuel will be sold for roughly $10/kg (O’Dell, 2015).

Yearly Revenue: $7.48 x 108

Simple Payback Period: 0.161 years

Return on Investment: 6.21

Profit after 10 years: $2.11 x 109

An NPV analysis was conducted to calculate the after-tax profits for the plant, taking into account variable depreciation as well as differing taxable income thresholds. The results can be seen below in Figure 2. Appendix E shows the calculation of the economic analysis.

[[File:hydrogenNPV.png|thumb|800px|center| Figure 2. NPV and Profit by year]]

==Optimization==

Heat exchangers were implemented where applicable to help reduce utility costs. The air flow into the heater was adjusted to the optimal level so that the heat necessary for the reaction was achieved and the savings from reduced fuel levels were greater than the additional cost of pumping air into the furnace. The temperature in the steam methane reformer was optimized to yield the quantity of hydrogen that provides the most revenue with respect to the amount of energy needed to heat the reformer. Optimization was also introduced by burning the unconverted gases that were separated as a result of PSA (Caventi et al. 2004). This saved money since the cost of recycling the unconverted gases back through the reformer again would be high due to the need for repressurization. The amine chosen for the process was also an area for optimization. Many amines will perform the job adequately, however, monoethanolamine (MEA) was chosen. MEA is cheap and since the process requires a large amount of purging, a cheap, effective amine is more viable than a more expensive one (Dow Chemical, 2003). Additionally, the sizes of the towers in the amine plant were increased to allow for more continuous operation, as it could operate for greater periods of time in between purges. Optimization allowed the required MEA feed to be reduced from 50000 kilomoles/hour to 9000 kilomoles/hour.

==Conclusion and Recommended Design Improvements==

With a simple payback period of 0.161 years and a return on investment of 6.21, the process is quite profitable and it is recommended to move forwards with the design. After conducting several rounds of optimization, however, it was determined that further improvements should be investigated to lower the required MEA in the amine plant and thereby reduce the major feedstock cost. This proposed plant has economic value and although the model only simulates the first 10 years of operation, the plant is expected to operate for 25 to 50 years given its capacity and robust design. This plant can sustain demand as the hydrogen market continues to grow in the United States and globally. Recommended design improvements include increasing the amount of amine recycled will lower feed costs, as this is one of the more expensive starting products utilized. Also, further optimization in the pressure swing adsorption to minimize costs, specifically the operational costs from the amount of energy utilized.

==References==

Actual Worldwide Hydrogen Production. 2007; http://www.fair-pr.de/background/worldwide-hydrogen-production-analysis.php. Accessed 1/15/2015, 2015.

Adris AM, Pruden BB. Canadian Journal of Chemical Engineering 74 (1996) 177

Alibaba. Coco fatty acid monoethanolamine CMEA 85% 97%. http://www.alibaba.com/product-detail/Coco-fatty-acid-monoethanolamine-CMEA-85_60179238235.html Accessed 3/3/2015.

Alibaba. Zibo Zeolite 3A, 4A, 5A, 13x Molecular Sieve in chemical adsorbent. http://www.alibaba.com/product-detail/Zibo-Zeolite-3A-4A-5A-13x_1183694288.html?s=p Accessed 3/3/2015

Baker RW. "Future Directions of Membrane Gas Separation Technology" Ind. Eng. Chem. Res. 2002, volume 41, pages 1393-1411.

Cavenati S, Grande C, Rodrigues A. Adsorption Equilibrium of Methane, Carbon Dioxide, and Nitrogen on Zeolite 13X at High Pressures. Journal of Chemical Engineering Data. 2004; 49(4):1095-1101.

Chandler BD, Schabel AB, Pignolet LH, Journal of Catalysis 193 (2000) 186–198.

ChE 351 Handout: Economic Evaluation of Projects: How Much is Our Idea Worth? 2014

Clay A. Boyce, MaCarr, Howe-Baker. Time for a New Hydrogen Plant? 2004. http://www.cbi.com/images/uploads/technical_articles/CBI_HydrocarbonEngineering_Feb04.pdf Accessed 1/13/2015.

Dow Chemical Company. Monoethanolamine, Diethanolamine, Triethanolamine. http://msdssearch.dow.com/PublishedLiteratureDOWCOM/dh_017d/0901b8038017d302.pdf?filepath=amines/pdfs/noreg/111-01375.pdf&fromPage=GetDoc 1/1/2003.

Freeport Water & Sewer Rates. City of Freeport, TX. 2014. Accessed 10/5/2014.

Holiday JD, Hu J, et al. An Overview of Hydrogen Production Technologies. Catalysis Today. 2007; 139(4):244-260.

Hoogers G. Fuel Cell Technology Handbook, CRC Press, Boca Raton (2003) pp. 5-
1–5-23.

Hydrogen Production – Steam Methane Reforming (SMR). 17 Columbia Circle, Albany, NY. 2004.

Knaebel, K. A “How To” Guide for Adsorber Design. http://userpages.umbc.edu/~dfrey1/ench445/AdsorberDes2.pdf Accessed 3/3/2015.

Kohl A, Nielson R. Gas Purification (5th ed.). Gulf Publishing (1997)

Miller LN, Zawacki TS. "Process for acid gas removal from gaseous mixtures", issued 21 Mar 1978, assigned to Institute of Gas Technology

O'Dell J. 8 Things You Need To Know About Hydrogen Fuel-Cell Cars. http://www.edmunds.com/fuel-economy/8-things-you-need-to-know-about-hydrogen-fuel-cell-cars.html. Accessed 3/2/2015.

Ogden JM. Review of Small Stationary Reformers for Hydrogen Production Princeton, NJ: Princeton University; 2002. http://www.afdc.energy.gov/pdfs/31948.pdf

Patt J, Moon DJ, Phillips C, Thompson L. Catalysis Letters 65 (2000) 193–195.

Rostrup-Nielsen J, in: I.T. Horvath (Ed.), Encyclopedia of Catalysis, Wiley Inter-science, (2003) 4.

Schoerner W, GS, Musich N. Selecting the Right Steam Methane Reformer: Can vs. Box Design. 2011. http://www.hydrocarbonprocessing.com/Article/2941986/Selecting-the-right- steam-methane-reformer-Can-vs-box-design.html

Song CS, Catalysis Today 77 (2002) 17–49

Song H, Zhang L, Ozkan US. Green Chemistry 9 (2007) 686–694

Steam Reforming. 2014. http://www.linde-engineering.com/en/process_plants /hydrogen and_ synthesis_gas_plants/gas_generation/steam_reforming/index.html. Accessed 1/13/2015.

Strasse, C.L. Hydrogen Recovery by Pressure Swing Adsorption. Linde Engineering 1/1/12. http://www.linde-engineering.com/internet.global.lindeengineering.global/en/images/HA_H_1_1_e_12_150dpi19_6130.pdf

TeGrotenhuis WE, King DL, Brooks KP, Golladay BJ, Wegeng RS, in: Baselt JP, Eul U, Wegeng RS (Eds.). Optimizing Microchannel Reactors by Trading- Off Equilibrium and Reaction Kinetics Through Temperature Management, AICHE, New Orleans, LA, 2002, p. 18.

Towler G, Sinnott R. Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013.

U.S. EIA. What is the average price of natural gas for electric-power-generation?. http://www.eia.gov/tools/faqs/faq.cfm?id=51&t=8 . Accessed 3/2/2015.

Wang HI; Koch WR, assignee. Patent 4869894. Hydrogen Generation and Recovery. 26 Sept. 1989. Print.

Wang Y, Chin Y, Rozmiarek RT, Watson J, Tonkovich ALY. Catalysis Today 98 (2004) 575–581.

What is PSA? Xebec Adsorption Inc. http://www.xebecinc.com/technology-what-is-psa.php. Accessed 1/30/15.

Yun S, Oyama S. Correlations in palladium membranes for hydrogen separation: A review. Journal of Membrane Science. 1989; 375(1-2):28-45.

==Appendix A: Stream and Vessel Number Descriptions==

'''Table 2 – Stream Number Descriptions'''
[[File:Table2streams.png|thumb|800px|center]]

Streams 100-117 are process cooling water streams for heat exchangers, the even numbered streams are the cooling water source and the odd numbered streams are the heated cooling water.
Stream 200 is process steam for heating in the reboiler and stream 201 is the steam condensate from amine reboiler.

'''Table 3 – Equipment Number Descriptions '''
[[File:Table3equip.PNG|thumb|800px|center]]

==Appendix B: Pressure Swing Adsorption Sizing ==
(Rostrup-Nielsen, 2003)

'''PSA Inlet Conditions'''
*5539 kmol/hour
*296.2 °C
*25 bar
*90.1% H2
*6.28% H2O
*1.82% CH4
*0.0026% CO2

'''Assumptions'''
*Cycle time: 10 min
*L/D = 4
*Bulk density of zeolite 5A: 660 kg/m3

5539 kmol/hr. = 2.147 x 104 kg/hr.

2.147×〖10〗^4 kg/hr((.468 kg H_2)/kg)((1 hr)/(6 cycles))(1/(.05 wt% load))(( 〖1 m〗^3)/(660 kg))(1/(0.85 (zeolite packing))) = 60 m3

Assume 2 sets of 3 columns

30/3 = 10 m3

10 + .1*10 = 11 m3 per column

Design Pressure: 25 bar * (1.1) = 27.5 bar = 400 psi

Design Temperature: 296 ºC + 50 ºF = 616 ºF

As L=4D

Volume = pi*D3 = 11 m3 → D=1.5 m

Hoop Stress: (0.4 ksi*1.52m)/(2*12.2*.85-1.2*.4ksi)=0.03 m ← highest stress

Long Stress: (0.4 ksi*1.52 m )/(4*12.2*0.85+0.8*.4 ksi)=0.15 m

Choose a hemispherical head for high pressure vessel

Hemispherical Head Stress: (0.4 ksi*1.52 m)/(4*12.2*.85-.4*0.4ksi)=0.0147 m

Surface area of sides: pi*D*L = 29.03 m2

Surface area of hemispheres: 4*pi*R2 = 7.26 m2

Total surface area: 36.29 m2

0.03 m x 36.29 m2 = 1.08 m3

1.08 m^3 ((7850 kg)/m^3 )=8547 kg carbon steel

Use Towler correlation for pressure vessel (p. 323)

C = a + bSn

11600 + 34(8547)0.85 = $103,000 per column

6 columns → $618,000

==Appendix C - Equipment Specifications==

'''Table 4 – Specifications for Process Equipment'''
[[File:Table4specs.PNG|thumb|800px|center]]

==Appendix D: Cost Estimation Details==

All calculations were made assuming the process will be operated for 24 hours a day, 360 days a year (to allow for holiday vacations).

Utilities: $29,817,760/yr. This value was obtained from the HYSYS economic analysis function.

Waste Treatment: $3,099,850/year at a cost of 1.5 dollars per metric ton (Towler and Sinnot, 2013). Calculated by summing up the total waste from the process on a yearly basis and multiplying by the cost per metric ton.

Labor: $6,750,000/year. This value was obtained using the following assumptions:
*One operator per operation unit paid $50,000/year
*Approximately 15 operation units (major units such as columns and reactors)
*3 eight-hour shifts per operation unit (covers 24 hours of operation/day)
*20% fudge factor on operator and supervisor wages to account for bonuses and overtime
*2 supervisors per eight-hour shift (1 for each portion of the plant)
*Supervisor pay is 25% of the total labor costs for operators.

Materials: $307,000,000/yr. Obtained by finding the yearly amount of water, methane, and MEA needed and multiplying by the price of water, methane, and MEA (Alibaba, 2015b).

Maintenance: $1,877,743/yr. This value was estimated as 3% of the capital cost (ChE 351 Handout, 2014).

Capital Investment: $62,591,460. This value was obtained from the HYSYS economic analysis function.

Yearly Revenue: $747,619,200/yr. Calculated by multiplying the cost of hydrogen per kilogram by the plant’s yearly output of hydrogen in kilograms (O’dell, 2015).

==Appendix E: Economic Calculations==

'''Simple Payback Period: '''

<math>SP=(initial cost)/(annual benefits-annual costs)</math>

<math>SP =62591460/(748000000-359000000) </math>

<math>SP = 0.161years</math>

'''Return on Investment: '''

<math>ROI= (annual benefits-annual costs)/(initial cost)</math>

<math> ROI =1/SP</math>

<math> ROI = 6.21</math>

'''Table 5: Net Present Value: '''

[[File: Table5npv.PNG|thumb|800px|center]]

Natural Gas to Hydrogen (H)

2015-03-13T20:19:45Z

Mzuckerman:

==Executive Summary==

==Introduction==

As the need to become less reliant on fossil fuels grows, the idea of widespread hydrogen fuel cells comes closer to being a reality. Hydrogen-fueled vehicles produce no direct pollution, making them environmentally friendly. However, fossil fuels are often used in the processes for the hydrogen production and the energy content of the produced hydrogen is less than that of the fossil fuels used (Clay et al. 2004; Actual Worldwide, 2007). To successfully implement a change to hydrogen fuel, high capacity plants must be constructed to meet increased demand. With the increased production of natural gas and the corresponding drop in price, steam reforming is a profitable method to produce hydrogen. The goal of this project is to design, optimize, and economically analyze a steam reforming process that will produce 99.999% hydrogen from a feedstock of natural gas. Waste products such as carbon dioxide, waste water, and monoethanolamine will be treated to minimize environmental impact.

__TOC__

==Technical Approach Taken==

After investigating process alternatives for each of the design stages, the final technologies chosen to achieve the specified design criteria were chosen by comparing the advantages and disadvantages as well as considering factors unique to the proposed design. The major decision was to use steam reforming to generate the hydrogen. The specific type of steam reformation chosen was conventional steam reforming due to capacity constraints. A nickel catalyst was chosen due to its low cost. To treat the CO waste, both methanation and a water gas shift were selected to ensure product purity. Amine scrubbing was picked to separate carbon dioxide from the product stream. Finally, pressure swing adsorption was chosen to handle separation of the methane from hydrogen. Several process design options were investigated to determine the most economically favorable and environmentally neutral design. The three major processes used for hydrogen production evaluated were steam reforming, partial oxidation, and autothermal reforming. Various technology options including reactor types, catalysts, and separation methods were compared to find the optimal combination of reliability, sustainability, profitability, and versatility.

The Peng-Robinson fluid package was chosen for use in the HYSYS simulation because it is ideal for describing hydrocarbon systems such as the steam methane reforming plant, where most components are gases or are nonpolar. Furthermore, Peng-Robinson is reliable over a large range of operating temperatures and pressures, which encompass those within our system. The amine plant portion of the simulation uses the amine package. This package is based on the Kent-Eisenberg model specifically designed for modeling removal of CO2 and H2S by amines.
Difficulties encountered modeling adsorption processes in HYSYS lead to the use of component splitters to model the feed pretreatment and pressure swing adsorption (PSA) steps.

==Steam Reforming Process==

The pre-treatment zone desulfurizes the hydrocarbon feed. Then, it is combined with superheated process steam. Afterwards, this mixture is heated and passes through reformer tubes filled with catalyst (Hydrogen Production, 2004). The reformer tubes are also heated which causes the hydrocarbon and steam mixture to react, yielding carbon monoxide, carbon dioxide, and hydrogen (Steam Reforming, 2014). The water gas shift then converts the carbon monoxide to carbon dioxide and generates hydrogen. The plant is normally operated with a higher steam content than required to prevent elemental carbon from being deposited in the catalyst bed (Steam Reforming, 2014). The pressure swing adsorption will isolate the hydrogen desired from any residual gases that are later used as fuel. Figure 1 below shows the process flow diagram of the process. The stream and vessel number descriptions are found in Appendix A.

==Process Flow Diagram==

[[File:hydrogenflowsheet.png|thumb|800px|center| Figure 1. Process Flow Diagram of the Hydrogen Reforming Process]]

==Process Alternatives==

Three primary techniques are used to produce hydrogen from a hydrocarbon source, such as methane. The techniques are: steam reforming, autothermal reforming, and partial oxidation.
Each of these techniques has advantages and disadvantages which must be considered in choosing a final design.

:;Steam reforming
::Advantages: industrially mature, no oxygen needed, lowest process operation temperature

::Disadvantage: large volume of air emissions

:;Autothermal reforming
::Advantage: low methane slip

::Disadvantages: requires air/oxygen, little commercial experience

:;Partial oxidation
::Advantages: no catalyst needed, needs less desulfurization, low methane slip

::Disadvantages: high process temperatures, process has high a degree of complexity, poor hydrogen to carbon monoxide ratio

Steam reforming is the most appropriate technique for the proposed design due to its high degree of industrial maturity and safe operating conditions. Although partial oxidation and autothermal reforming each have distinct advantages, the lack of maturity for both of these processes would add a high degree of risk to the overall plant design. Furthermore, the lower operation temperature of steam reforming will result in a safer process (Holiday et al. 2007).

The five methods of steam reforming that would provide the necessary purity for hydrogen cars are the conventional box method, conventional can method, the compact steam methane reformers, plate-type steam methane reformers, and membrane reactor steam reforming.

:;Conventional box steam reforming
::Advantages: Conventional reforming can handle a large throughput such as the 100 MMscfd. The desired 99.99% purity is achieved after the water-gas shift and pressure swing adsorption (Ogden, 2002).6

::Disadvantages: Temperatures of 750-800 °C are needed in the reactor. The utility costs associated with this method are high (Schoerner et al. 2011).7

:;Conventional can steam reforming
::Advantages: Can be quickly constructed for fast start up.

::Disadvantages: Only cost efficient at throughputs under 10 MMscfd so multiple would need to be used (Schoerner et al. 2011).

:;Compact steam methane reformers
::Advantages: Concentric annular catalyst beds can produce the 99.999% purity which is more pure than required.

::Disadvantages: Produce 1 MMscfd and, therefore, this alternative is unable to handle the throughput of the plant being designed (Ogden, 2002).

:;Plate-type steam methane reformers
::Advantages: Plate-type steam reformers are less expensive, more efficient, and have faster startups than conventional methods.

::Disadvantages: The catalysts used are extremely sensitive to carbon deposition and sulfur poisoning. Implementing this process is more risky as it is not yet being used commercially (Ogden, 2002).

:;Membrane reactors for steam reforming
::Advantages: This method produces extremely pure product.

::Disadvantages: Membrane reactors require temperatures 200 °C greater than conventional methods and does not work effectively at the large scale desired (Ogden, 2002).

After taking all the processes into consideration, the conventional box steam reformer is the most appropriate choice for the plant being designed as it will yield the desired purity and can produce the 100 MMscfd desired.

==Technology Options==

===Reactor and Catalyst Alternatives===

Of the several process alternatives described above, steam reforming is the most developed technology for hydrogen production. However, there are different ways to carry out the process of steam reforming. The alternatives focus on the choice of catalyst and the configuration of the catalyst in the reactor.

Generally, catalysts can be categorized as non-precious metal (such as nickel) and precious metal (such as platinum and rhodium). The cost of non-precious metal catalysts is substantially lower, but the catalyst is also far less effective. Nevertheless, conventional steam reforming plants usually use nickel catalysts because heat and mass transfer effects generally dominate the reaction kinetics with an effectiveness factor as low as 5% (Adris and Pruden, 1996).

To reduce the heat and mass transfer effects, some processes attempt to reduce the particle size of catalysts and increase the reaction area by using microchannel-based reactors (Wang et al. 2004). As the size of the catalyst particle decreases, the apparent kinetics of reaction converge to the intrinsic kinetic of the catalyst. As a result, the more expensive precious metal catalysts are favored (Rostrup-Nielsen, 2003). Currently, researchers are developing cobalt-based catalysts to mitigate the cost of rhodium based catalysts (Song et al. 2007).

===Carbon Monoxide Removal Alternatives===

The steam reforming process produces approximately 5% of carbon monoxide in the product stream. Water-gas shift and methanation are two most common method to improve purity (Song, 2002). Water-gas shift reacts carbon monoxide and water to form hydrogen and carbon dioxide. In order to prevent the oxidation of hydrogen, precise control of air input is needed. Conventionally, the product stream is passed through a high temperature reactor because the reaction is substantially faster at higher temperature, but it is then passed through a low temperature reactor to improve hydrogen production (Hoogers, 2003). TeGrotenhuis et al. (2002) have demonstrated that a single reactor with a gradient temperature that spans both ends of the temperature extremes can be used in order to reduce the utility used. While copper catalysts are most prevalent, increased selectivity can be achieved with higher cost molybdenum carbide or platinum-based catalysts (Patt et al. 2000; Chandler et al. 200).

Another process to reduce carbon monoxide is methanation. Similar to water-gas shift, methanation reacts carbon monoxide and hydrogen to form methane and water. Methanation reactors are generally simpler without the need for a precise air stream. However, the reaction fundamentally consumes hydrogen at a very high rate (3 moles of hydrogen are consumed for 1 mole of carbon monoxide consumed) (Hoogers, 2003). A combination of both methods is generally used for high purity hydrogen production, which is appropriate for the proposed design. As such, both a water-gas shift as well as methanation will be utilized.

===Acid Gas Removal===

After the water gas shift, the product stream will be enriched in carbon dioxide and methane in order to remove carbon monoxide. The two stack gases will be separated in two steps. Carbon dioxide is an acid gas, which is generally removed by adsorption using an amine solvent (Kohl and Nielson, 1997). The product stream is passed through an absorber with the amine solvent where the carbon dioxide is stripped. The amine, now rich in carbon dioxide is then passed through a regenerator where the acid gas is separated from the amine (Miller and Zawacki, 1978). This process is referred to as amine scrubbing.

Other technologies that are used for acid gas removal include using a membrane to separate the acid gas and vacuum swing separation. Membrane gas separation however is still not mature enough for widespread usage today (Baker, 2002). Another alternative is to use vacuum swing adsorption. The feed gas is passed through a parallel adsorption bed where the carbon dioxide is adsorbed. The bed is then rinsed with a stream rich in carbon dioxide to remove any methane or hydrogen trapped in the bed. The pressure within the adsorber is then lowered to further the separation of the gases (Wang, 1989). Due to the complexity of the vacuum swing adsorption, and the immaturity of membrane separation, using an amine solvent for amine scrubbing is the method of choice for removing the carbon dioxide.

===Methane Removal===

Two potential methods of separating methane from hydrogen and water are membrane gas separation and pressure swing adsorption. Membrane gas separation is the less mature and reliable of these technologies. It involves the use of either a porous or non-porous membrane in order to selectively separate gas components primarily on size and affinity. Synthetic membranes are contained in either hollow fiber bundles, spiral wound bundles, or plates (Yun and Oyama, 1989).

Pressure swing adsorption (PSA) selectively adsorbs gas components by rapidly pressurizing and depressurizing a gas around adsorbents. Species are adsorbed at high pressure and the adsorbents can be regenerated by and series of purging and desorbing at low pressures (What is PSA, 2015). Common adsorbent materials are zeolites, molecular sieves, silica gel, alumina, and activated carbon. PSA operates at near-ambient temperatures and is the most widely accepted method of large scale hydrogen purification (Strasse, 2012).

==Key Processes and Products==

There are seven major processing steps involved in the steam methane reformation process as designed. These major processing stages are: the initial heating of feed, the steam reformation reactor, the high and low temperature water shift reactors, the amine plant, the methanation reactor, the gas compression and cooling train, and the pressure swing adsorption unit.

From the heating of the feed, the combustion of hydrocarbons to provide heat yields carbon dioxide as a major byproduct. The product of this process is the heated feed which contains methane and water. The main product of the steam reformation reactor contains methane, carbon dioxide, carbon monoxide, water, and methane. This product goes to the shift reactors where the outgoing product will contain less carbon monoxide and more carbon dioxide than that from the steam reformation reactor, but will otherwise contain similar components. The product from the shift reactors is sent to an amine plant where the acid gas (Carbon Dioxide) will be removed and vented to the atmosphere as waste product, resulting in a product stream containing carbon monoxide, methane, hydrogen, and water. This stream is sent to the methanation reactor where the carbon monoxide is broken down into oxygen and carbon which yields a product stream containing water, methane, and hydrogen. This stream is then sent to the compression system where the pressure is increased so that pressure swing adsorption may be used to separate the methane and hydrogen. From the pressure swing adsorption unit, the purified hydrogen is removed and sent to a storage tank to be sold later and the methane captured is then recycled back to the gas fired heater. For more details on the products in each flow, see Appendix A.

==Equipment Sizing==

Two sources were used to size the equipment and estimate capital costs. The first was the Aspen HYSYS integrated economics evaluator. Carbon steel was the chosen material for use in the furnace, pumps, and pipes. All components were sized with a 10% pressure safety margin or 25 psi greater than the operating conditions (whichever value was higher) in order to provide a potential buffer in case of an increase in temperature or pressure. This resulted in moderately higher prices, but no expense was spared due to safety concerns. The reactors were sized using methods described in Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design.

Since the PSA columns could not be modeled in HYSYS, individual suppliers were contacted to achieve the required plant capacity. Using empirical correlations, the capital cost was estimated for six PSA columns, shown in Appendix B. The equipment sizes (diameter, etc.) for certain equipment utilized in the process are shown in Appendix C.

==Economic Analysis==

Total Capital Cost: $6.26 x 107

Table 1 shows the annual operating costs for the process. More detail on how these costs were calculated can be found in Appendix D.

{| class="wikitable" style="margin: auto;"
!colspan="2"| Table 1: Annual Operating Costs
|-
! Process Expense
! Yearly Cost ($)
|-
| Waste Treatment
|1.86 x 106
|-
| Utilities
|2.98 x 107
|-
| Labor
|1.80 x 107
|-
| Maintenance
|1.88 x 106
|-
| Feedstock Purchase
|3.07 x 108
|}

Total Annual Operating Cost: $3.59 x 108, assuming the process will be operated for 24 hours a day, 360 days a year (to allow for holiday vacations).

The water will be sold at $0.00207 per kilogram (Freeport Water, 2014). Methane will be purchased for $0.226 per kilogram (U.S. EIA, 2015). MEA will be purchased for $0.50 per kilogram (Alibaba, 2015a). Hydrogen fuel will be sold for roughly $10/kg (O’Dell, 2015).

Yearly Revenue: $7.48 x 108

Simple Payback Period: 0.161 years

Return on Investment: 6.21

Profit after 10 years: $2.11 x 109

An NPV analysis was conducted to calculate the after-tax profits for the plant, taking into account variable depreciation as well as differing taxable income thresholds. The results can be seen below in Figure 2. Appendix E shows the calculation of the economic analysis.

[[File:hydrogenNPV.png|thumb|800px|center| Figure 2. NPV and Profit by year]]

==Optimization==

Heat exchangers were implemented where applicable to help reduce utility costs. The air flow into the heater was adjusted to the optimal level so that the heat necessary for the reaction was achieved and the savings from reduced fuel levels were greater than the additional cost of pumping air into the furnace. The temperature in the steam methane reformer was optimized to yield the quantity of hydrogen that provides the most revenue with respect to the amount of energy needed to heat the reformer. Optimization was also introduced by burning the unconverted gases that were separated as a result of PSA (Caventi et al. 2004). This saved money since the cost of recycling the unconverted gases back through the reformer again would be high due to the need for repressurization. The amine chosen for the process was also an area for optimization. Many amines will perform the job adequately, however, monoethanolamine (MEA) was chosen. MEA is cheap and since the process requires a large amount of purging, a cheap, effective amine is more viable than a more expensive one (Dow Chemical, 2003). Additionally, the sizes of the towers in the amine plant were increased to allow for more continuous operation, as it could operate for greater periods of time in between purges. Optimization allowed the required MEA feed to be reduced from 50000 kilomoles/hour to 9000 kilomoles/hour.

==Conclusion and Recommended Design Improvements==

With a simple payback period of 0.161 years and a return on investment of 6.21, the process is quite profitable and it is recommended to move forwards with the design. After conducting several rounds of optimization, however, it was determined that further improvements should be investigated to lower the required MEA in the amine plant and thereby reduce the major feedstock cost. This proposed plant has economic value and although the model only simulates the first 10 years of operation, the plant is expected to operate for 25 to 50 years given its capacity and robust design. This plant can sustain demand as the hydrogen market continues to grow in the United States and globally. Recommended design improvements include increasing the amount of amine recycled will lower feed costs, as this is one of the more expensive starting products utilized. Also, further optimization in the pressure swing adsorption to minimize costs, specifically the operational costs from the amount of energy utilized.

==References==

Actual Worldwide Hydrogen Production. 2007; http://www.fair-pr.de/background/worldwide-hydrogen-production-analysis.php. Accessed 1/15/2015, 2015.

Adris AM, Pruden BB. Canadian Journal of Chemical Engineering 74 (1996) 177

Alibaba. Coco fatty acid monoethanolamine CMEA 85% 97%. http://www.alibaba.com/product-detail/Coco-fatty-acid-monoethanolamine-CMEA-85_60179238235.html Accessed 3/3/2015.

Alibaba. Zibo Zeolite 3A, 4A, 5A, 13x Molecular Sieve in chemical adsorbent. http://www.alibaba.com/product-detail/Zibo-Zeolite-3A-4A-5A-13x_1183694288.html?s=p Accessed 3/3/2015

Baker RW. "Future Directions of Membrane Gas Separation Technology" Ind. Eng. Chem. Res. 2002, volume 41, pages 1393-1411.

Cavenati S, Grande C, Rodrigues A. Adsorption Equilibrium of Methane, Carbon Dioxide, and Nitrogen on Zeolite 13X at High Pressures. Journal of Chemical Engineering Data. 2004; 49(4):1095-1101.

Chandler BD, Schabel AB, Pignolet LH, Journal of Catalysis 193 (2000) 186–198.

ChE 351 Handout: Economic Evaluation of Projects: How Much is Our Idea Worth? 2014

Clay A. Boyce, MaCarr, Howe-Baker. Time for a New Hydrogen Plant? 2004. http://www.cbi.com/images/uploads/technical_articles/CBI_HydrocarbonEngineering_Feb04.pdf Accessed 1/13/2015.

Dow Chemical Company. Monoethanolamine, Diethanolamine, Triethanolamine. http://msdssearch.dow.com/PublishedLiteratureDOWCOM/dh_017d/0901b8038017d302.pdf?filepath=amines/pdfs/noreg/111-01375.pdf&fromPage=GetDoc 1/1/2003.

Freeport Water & Sewer Rates. City of Freeport, TX. 2014. Accessed 10/5/2014.

Holiday JD, Hu J, et al. An Overview of Hydrogen Production Technologies. Catalysis Today. 2007; 139(4):244-260.

Hoogers G. Fuel Cell Technology Handbook, CRC Press, Boca Raton (2003) pp. 5-
1–5-23.

Hydrogen Production – Steam Methane Reforming (SMR). 17 Columbia Circle, Albany, NY. 2004.

Knaebel, K. A “How To” Guide for Adsorber Design. http://userpages.umbc.edu/~dfrey1/ench445/AdsorberDes2.pdf Accessed 3/3/2015.

Kohl A, Nielson R. Gas Purification (5th ed.). Gulf Publishing (1997)

Miller LN, Zawacki TS. "Process for acid gas removal from gaseous mixtures", issued 21 Mar 1978, assigned to Institute of Gas Technology

O'Dell J. 8 Things You Need To Know About Hydrogen Fuel-Cell Cars. http://www.edmunds.com/fuel-economy/8-things-you-need-to-know-about-hydrogen-fuel-cell-cars.html. Accessed 3/2/2015.

Ogden JM. Review of Small Stationary Reformers for Hydrogen Production Princeton, NJ: Princeton University; 2002. http://www.afdc.energy.gov/pdfs/31948.pdf

Patt J, Moon DJ, Phillips C, Thompson L. Catalysis Letters 65 (2000) 193–195.

Rostrup-Nielsen J, in: I.T. Horvath (Ed.), Encyclopedia of Catalysis, Wiley Inter-science, (2003) 4.

Schoerner W, GS, Musich N. Selecting the Right Steam Methane Reformer: Can vs. Box Design. 2011. http://www.hydrocarbonprocessing.com/Article/2941986/Selecting-the-right- steam-methane-reformer-Can-vs-box-design.html

Song CS, Catalysis Today 77 (2002) 17–49

Song H, Zhang L, Ozkan US. Green Chemistry 9 (2007) 686–694

Steam Reforming. 2014. http://www.linde-engineering.com/en/process_plants /hydrogen and_ synthesis_gas_plants/gas_generation/steam_reforming/index.html. Accessed 1/13/2015.

Strasse, C.L. Hydrogen Recovery by Pressure Swing Adsorption. Linde Engineering 1/1/12. http://www.linde-engineering.com/internet.global.lindeengineering.global/en/images/HA_H_1_1_e_12_150dpi19_6130.pdf

TeGrotenhuis WE, King DL, Brooks KP, Golladay BJ, Wegeng RS, in: Baselt JP, Eul U, Wegeng RS (Eds.). Optimizing Microchannel Reactors by Trading- Off Equilibrium and Reaction Kinetics Through Temperature Management, AICHE, New Orleans, LA, 2002, p. 18.

Towler G, Sinnott R. Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013.

U.S. EIA. What is the average price of natural gas for electric-power-generation?. http://www.eia.gov/tools/faqs/faq.cfm?id=51&t=8 . Accessed 3/2/2015.

Wang HI; Koch WR, assignee. Patent 4869894. Hydrogen Generation and Recovery. 26 Sept. 1989. Print.

Wang Y, Chin Y, Rozmiarek RT, Watson J, Tonkovich ALY. Catalysis Today 98 (2004) 575–581.

What is PSA? Xebec Adsorption Inc. http://www.xebecinc.com/technology-what-is-psa.php. Accessed 1/30/15.

Yun S, Oyama S. Correlations in palladium membranes for hydrogen separation: A review. Journal of Membrane Science. 1989; 375(1-2):28-45.

==Appendix A: Stream and Vessel Number Descriptions==

'''Table 2 – Stream Number Descriptions'''
[[File:Table2streams.png|thumb|800px|center]]

Streams 100-117 are process cooling water streams for heat exchangers, the even numbered streams are the cooling water source and the odd numbered streams are the heated cooling water.
Stream 200 is process steam for heating in the reboiler and stream 201 is the steam condensate from amine reboiler.

'''Table 3 – Equipment Number Descriptions '''
[[File:Table3equip.PNG|thumb|800px|center]]

==Appendix B: Pressure Swing Adsorption Sizing ==
(Rostrup-Nielsen, 2003)

'''PSA Inlet Conditions'''
*5539 kmol/hour
*296.2 °C
*25 bar
*90.1% H2
*6.28% H2O
*1.82% CH4
*0.0026% CO2

'''Assumptions'''
*Cycle time: 10 min
*L/D = 4
*Bulk density of zeolite 5A: 660 kg/m3

5539 kmol/hr. = 2.147 x 104 kg/hr.

2.147×〖10〗^4 kg/hr((.468 kg H_2)/kg)((1 hr)/(6 cycles))(1/(.05 wt% load))(( 〖1 m〗^3)/(660 kg))(1/(0.85 (zeolite packing))) = 60 m3

Assume 2 sets of 3 columns

30/3 = 10 m3

10 + .1*10 = 11 m3 per column

Design Pressure: 25 bar * (1.1) = 27.5 bar = 400 psi

Design Temperature: 296 ºC + 50 ºF = 616 ºF

As L=4D

Volume = pi*D3 = 11 m3 → D=1.5 m

Hoop Stress: (0.4 ksi*1.52m)/(2*12.2*.85-1.2*.4ksi)=0.03 m ← highest stress

Long Stress: (0.4 ksi*1.52 m )/(4*12.2*0.85+0.8*.4 ksi)=0.15 m

Choose a hemispherical head for high pressure vessel

Hemispherical Head Stress: (0.4 ksi*1.52 m)/(4*12.2*.85-.4*0.4ksi)=0.0147 m

Surface area of sides: pi*D*L = 29.03 m2

Surface area of hemispheres: 4*pi*R2 = 7.26 m2

Total surface area: 36.29 m2

0.03 m x 36.29 m2 = 1.08 m3

1.08 m^3 ((7850 kg)/m^3 )=8547 kg carbon steel

Use Towler correlation for pressure vessel (p. 323)

C = a + bSn

11600 + 34(8547)0.85 = $103,000 per column

6 columns → $618,000

==Appendix C - Equipment Specifications==
'''Table 4 – Specifications for Process Equipment'''
[[File:Table4specs.PNG|thumb|800px|center]]

==Appendix D: Cost Estimation Details==

All calculations were made assuming the process will be operated for 24 hours a day, 360 days a year (to allow for holiday vacations).

Utilities: $29,817,760/yr. This value was obtained from the HYSYS economic analysis function.

Waste Treatment: $3,099,850/year at a cost of 1.5 dollars per metric ton (Towler and Sinnot, 2013). Calculated by summing up the total waste from the process on a yearly basis and multiplying by the cost per metric ton.

Labor: $6,750,000/year. This value was obtained using the following assumptions:
*One operator per operation unit paid $50,000/year
*Approximately 15 operation units (major units such as columns and reactors)
*3 eight-hour shifts per operation unit (covers 24 hours of operation/day)
*20% fudge factor on operator and supervisor wages to account for bonuses and overtime
*2 supervisors per eight-hour shift (1 for each portion of the plant)
*Supervisor pay is 25% of the total labor costs for operators.

Materials: $307,000,000/yr. Obtained by finding the yearly amount of water, methane, and MEA needed and multiplying by the price of water, methane, and MEA (Alibaba, 2015b).

Maintenance: $1,877,743/yr. This value was estimated as 3% of the capital cost (ChE 351 Handout, 2014).

Capital Investment: $62,591,460. This value was obtained from the HYSYS economic analysis function.

Yearly Revenue: $747,619,200/yr. Calculated by multiplying the cost of hydrogen per kilogram by the plant’s yearly output of hydrogen in kilograms (O’dell, 2015).

==Appendix E: Economic Calculations==

'''Simple Payback Period: '''

<math>SP=(initial cost)/(annual benefits-annual costs)</math>

<math>SP =62591460/(748000000-359000000) </math>

<math>SP = 0.161years</math>

'''Return on Investment: '''

<math>ROI= (annual benefits-annual costs)/(initial cost)</math>

<math> ROI =1/SP</math>

<math> ROI = 6.21</math>

'''Table 5: Net Present Value: '''

[[File: Table5npv.PNG|thumb|800px|center]]

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Natural Gas to Hydrogen (H)

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Mzuckerman: Created page with "==Executive Summary== ==Introduction== As the need to become less reliant on fossil fuels grows, the idea of widespread hydrogen ..."

Pinch analysis

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Mzuckerman:

Author: Micah Zuckerman [2015] 

Stewards: Jian Gong and Fengqi You

__TOC__

==Introduction==
Pinch analysis is a methodology used to minimize energy consumption of chemical processes by optimizing energy recovery methods. A pinch occurs at the closest point of approach between hot and cold streams in the network. By designing around this point energy from utilities can be minimized. This page focuses on pinch analysis in the context of heat exchanger networks (HENs), however similar methods can be used in the design of mass exchanger networks (MENs).
==Pinch technology==
For maximum energy recovery (MER), energy should not be transferred across the pinch. The network is separated into hot and cold regions, above and below the pinch respectively. In the hot region, streams are matched so that heat is transferred from the coolest possible source. Below the pinch, the opposite occurs. This yields the minimum utility requirement solution to the heat exchanger network.
Minimum utility requirement is not necessarily the optimum solution, however, the solution with the lowest overall cost is. Minimizing the size and quantity of heat exchangers plays a large role in decreasing capital costs. There is a trade-off between operational and capital costs. In some cases transfer across the pinch is necessary to produce the optimal network. For more information see the trade-off section of the [[Sensitivity analysis and design optimization]] page.

==Minimum approach temperature==
The minimum approach temperature represents the smallest gap in temperature across which heat transfer will occur in the system. This often only occurs at the pinch. While technically any value greater than zero can allow for heat transfer, very small values are not often feasible. Low temperature differences decrease the need for additional utilities, but require increasingly large heat transfer areas, meaning larger heat exchangers. The trade-off between operating costs and capital costs should be used to choose the minimum approach temperature for the network. Typical choices for minimum approach temperatures are between 5 °C and 30 °C (Towler and Sinnott 2013).
==Analysis Methods==
Pinch technology is a highly used and advanced technology throughout industry. Because of this, there have been devised multiple approaches with which to solve a pinch system.
===Composite curves method===
[[File:Composite Curves example.png|thumb|600px|center|Example of composite curve method taken from Sieder et al. Fig 10.7 ]]
This method uses a graphical approach and makes obvious the presence and location of the pinch in the system. First one combines all hot streams into single composite curve on a temperature-enthalpy graph. The same is then done with the cold streams. Plotting both composite curves on a temperature-enthalpy plot reveals a pinch at the point where the cold and hot curves are closest in temperature. Because change in enthalpy is the feature of concern, the composite curves can be freely shifted along the enthalpy axis. In this manner the minimum approach temperature can be satisfied at the pinch.
===Problem table method===
The problem table method, also known as the cascade method, forgoes graphs for a numerical approach. This method is useful when solving heat-exchanger network problems manually, but has fallen into disuse due to the advent of modern computing. The first step in the process is converting stream temperatures into interval temperatures. To do this, one subtracts half the minimum approach temperature from the hot stream temperatures, and adds half to the cold stream temperatures. The intervals are then ranked in order of magnitude, ignoring any repetitions. The next step involves balancing the heat carried by streams between each interval. Any excess heat cascades over into the next lower down interval. Add heat from the top of the system eliminate any negative values. The amount added at the top corresponds to hot utility duty, while the remaining value at the bottom corresponds to heat that must be removed by cold utility.
[[File:Cascade.PNG|thumb|500px|center|Example of the problem table method. The right most column contains both the minimum heating requirement for the system at the top, and the minimum cooling requirement at the bottom]]
===Computer tools for HEN design===
Presently, the two previously described methods of pinch analysis are not in wide use in industry; rather analysis is performed by specialized software. These software include: Aspen HX-Net™ from Aspen Technology Inc., SUPERTARGET™ from Linnhoff March Ltd. and UniSim™ ExchangerNet™ from Honeywell International Inc. (Towler and Sinnott 2013). Stream data is often pulled directly from the process simulations by the software package. Automatically generated networks must be properly constrained lest an unfeasible amount of stream splits and loops be generated. There is great value, however, in the ability of software packages to perform many tweaks and iterations to fine tune a network.
==Example problem==
In this example problem we will use pinch analysis via the problem table method to design a heat exchanger network given the following stream table, taken from Towler and Sinnott, and a desired minimum approach temperature of 20°C.

{|class="wikitable" style="margin: auto;"

! Stream
! Number Type
! Heat Capacity Flow Rate, kW/°C
! Source Temp. °C
! Target Temp. °C
! Heat Load, kW
|-
|1
| hot
| 40.0
| 180
| 40
| 5600
|-
|2
| hot
| 30.0
|150
| 60
| 1800
|-
|3
| cold
| 60.0
| 30
| 180
| 9000
|-
|4
| cold
| 20.0
| 80
| 160
| 1600
|}

The first step is to calculate the interval temperatures. Because the minimum approach temperature is 20°C, the interval temperatures for the hot stream are 10 °C less than the actual temperature and those for the cold stream are 10°C more than the actual temperature. The next step is to rank the interval temperatures, ignoring any duplicated values. We then determine what streams belong in which interval. The sum of CP values for the cold streams minus that for the hot streams can then be multiplied by the
interval ΔT to give the interval ΔH, and the interval ΔH values can be cascaded to give the overall heat flow.

[[File: Problem table example.PNG|thumb|500px|center|Problem table exhibiting cascade taken from Figure 3.28 of Towler and Sinnott]]

The amount of heat added at the top to remove any negative values, 2900 kW, is the hot utility. The 600 kW left over at the bottom is the cold utility required. The pinch occurs where the heat transferred is zero, that is at interval number 4, interval temperature 90 °C. This is equivalent to 100°C in the hot stream or 80°C in the cold stream.

Now we need to design the actual system:

'''Above Pinch'''
# CPh ≤ CPc
# It is possible to match stream 1 or 2 with stream 3 but neither stream can match with stream 4. This is problematic, since if we match stream 1 with 3 then stream 2 will not be able to make a match at the pinch. Likewise, if we match stream 2 with 3 then stream 1 will not have a match at the pinch.
# Check the heat available in bringing the hot streams to the pinch temperature.
#::stream 1 ΔH = 40.0(180 − 100 )= 3200 kW
#::stream 2 ΔH = 30.0 (150 − 100 ) = 1500 kW
# Check the heat required to bring the cold streams from the pinch temperature to their target temperatures.
#::stream 3 ΔH = 60.0(180 – 80)= 6000 kW
#::stream 4 ΔH = 20.0(160 – 80) = 1600 kW
# If we split stream 3 into two branches with CP of 40.0 and 20.0, then we can match the larger branch with stream 1 and transfer 3200 kW, which satisfies (ticks off) stream 1.
# We now have two cold streams, both with CP of 20.0, and one hot stream (2) with CP of 30.0. We need to split stream 2 into two branches. As an initial guess these can both have CP of 15.0. We can then match one branch of stream 2 with the smaller branch of 4 and transfer 750 kW, and the other branch with stream 3, also for 750 kW, which then ticks off stream 2.
# Include a heater on the larger branch of stream 3 to bring it to its target temperature:
#::ΔHhot = 40 (100 ) − 3200 = 800 kW
# Include a heater on the smaller branch of stream 3 to provide the balance of the heat required:
#::ΔHhot = 20 (100 ) – 750 = 1250 kW
# Include a heater on stream 4 to provide the balance of the heat required:
#::Δ Hhot = 1600 − 750 = 850 kW
# Check sum of heater duties = 800 + 1250 + 850 = 2900 kW = hot utility target.
'''Below Pinch'''
# CPh ≥ CPc
# Note that stream 4 starts at the pinch temperature and so cannot provide any cooling below the pinch.
# We cannot match stream 1 or 2 with stream 3 at the pinch.
# Split stream 3 to reduce CP. An even split will allow both streams 1 and 2 to be matched with the split streams adjacent to the pinch, so try this initially.
# Check the heat available from bringing the hot streams from the pinch temperature to their target temperatures:
#::stream 1 ΔH = 40.0 (100 − 40 ) = 2400 kW
#::stream 2 ΔH = 30.0 (100 − 60 ) = 1200 kW
# Check the heat required to bring the cold streams from their source temperatures to the pinch temperature:
#::stream 3 ΔH = 60.0 (80 − 30 ) = 3000 kW
#::Stream 4 is at the pinch temperature
# Note that stream 1 cannot be brought to its target temperature of 40 °C by full interchange with stream 3 as the source temperature of stream 3 is 30 °C, so ΔTmin would be violated. So transfer 1800 kW to one leg of the split stream 3.
# Check temperature at exit of this exchanger:
#::Temp out= 100 – 1800/40= 55 °C, satisfactory
# Provide cooler on stream 1 to bring it to its target temperature; the cooling needed is
#::ΔHcold = 2400 − 1800 = 600 kW
# Transfer the full heat load from stream 2 to second leg of stream 3; this satisfies both streams.
[[File: Pinch example.PNG|thumb|500px|center|Proposed Heat Exchanger Network for example problem from Figure 3.29 of Towler and Sinnott]]

==References==
Biegler LT, Grossmann IE, Westerberg AW. Systematic Methods of Chemical Process Design. Upper Saddle River: Prentice Hall; 1997.

Seider WD, Seader JD, Lewin DR. Process Design Principles: Synthesis, Analysis, and Evaluation. 3rd ed. New York: Wiley; 2004.

Towler G, Sinnott R. Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013.

Turton R, Bailie RC, Whiting WB, Shaewitz JA, Bhattacharyya D. Analysis, Synthesis, and Design of Chemical Processes. 4th ed. Upper Saddle River: Prentice-Hall; 2012.

File:Problem table example.PNG

2015-03-02T05:50:00Z

Mzuckerman:

File:Pinch example.PNG

2015-03-02T03:48:26Z

Mzuckerman:

Pinch analysis

2015-03-02T03:47:29Z

Mzuckerman:

Author: Micah Zuckerman [2015] 

Stewards: Jian Gong and Fengqi You

__TOC__

==Introduction==
Pinch analysis is a methodology used to minimize energy consumption of chemical processes by optimizing energy recovery methods. A pinch occurs at the closest point of approach between hot and cold streams in the network. By designing around this point energy from utilities can be minimized. This page focuses on pinch analysis in the context of heat exchanger networks (HENs), however similar methods can be used in the design of mass exchanger networks (MENs).
==Pinch technology==
For maximum energy recovery (MER), energy should not be transferred across the pinch. The network is separated into hot and cold regions, above and below the pinch respectively. In the hot region, streams are matched so that heat is transferred from the coolest possible source. Below the pinch, the opposite occurs. This yields the minimum utility requirement solution to the heat exchanger network.
Minimum utility requirement is not necessarily the optimum solution, however, the solution with the lowest overall cost is. Minimizing the size and quantity of heat exchangers plays a large role in decreasing capital costs. There is a trade-off between operational and capital costs. In some cases transfer across the pinch is necessary to produce the optimal network. For more information see the trade-off section of the [[Sensitivity analysis and design optimization]] page.

==Minimum approach temperature==
The minimum approach temperature represents the smallest gap in temperature across which heat transfer will occur in the system. This often only occurs at the pinch. While technically any value greater than zero can allow for heat transfer, very small values are not often feasible. Low temperature differences decrease the need for additional utilities, but require increasingly large heat transfer areas, meaning larger heat exchangers. The trade-off between operating costs and capital costs should be used to choose the minimum approach temperature for the network. Typical choices for minimum approach temperatures are between 5 °C and 30 °C (Towler and Sinnott 2013).
==Analysis Methods==
Pinch technology is a highly used and advanced technology throughout industry. Because of this, there have been devised multiple approaches with which to solve a pinch system.
===Composite curves method===
[[File:Composite Curves example.png|thumb|600px|center|Example of composite curve method taken from Sieder et al. Fig 10.7 ]]
This method uses a graphical approach and makes obvious the presence and location of the pinch in the system. First one combines all hot streams into single composite curve on a temperature-enthalpy graph. The same is then done with the cold streams. Plotting both composite curves on a temperature-enthalpy plot reveals a pinch at the point where the cold and hot curves are closest in temperature. Because change in enthalpy is the feature of concern, the composite curves can be freely shifted along the enthalpy axis. In this manner the minimum approach temperature can be satisfied at the pinch.
===Problem table method===
The problem table method, also known as the cascade method, forgoes graphs for a numerical approach. This method is useful when solving heat-exchanger network problems manually, but has fallen into disuse due to the advent of modern computing. The first step in the process is converting stream temperatures into interval temperatures. To do this, one subtracts half the minimum approach temperature from the hot stream temperatures, and adds half to the cold stream temperatures. The intervals are then ranked in order of magnitude, ignoring any repetitions. The next step involves balancing the heat carried by streams between each interval. Any excess heat cascades over into the next lower down interval. Add heat from the top of the system eliminate any negative values. The amount added at the top corresponds to hot utility duty, while the remaining value at the bottom corresponds to heat that must be removed by cold utility.
[[File:Cascade.PNG|thumb|500px|center|Example of the problem table method. The right most column contains both the minimum heating requirement for the system at the top, and the minimum cooling requirement at the bottom]]
===Computer tools for HEN design===
Presently, the two previously described methods of pinch analysis are not in wide use in industry; rather analysis is performed by specialized software. These software include: Aspen HX-Net™ from Aspen Technology Inc., SUPERTARGET™ from Linnhoff March Ltd. and UniSim™ ExchangerNet™ from Honeywell International Inc. (Towler and Sinnott 2013). Stream data is often pulled directly from the process simulations by the software package. Automatically generated networks must be properly constrained lest an unfeasible amount of stream splits and loops be generated. There is great value, however, in the ability of software packages to perform many tweaks and iterations to fine tune a network.
==Example==

[[File:Pinch example.PNG|thumb|500px|center|Proposed Heat Exchanger Network for example problem from Figure 3.29 of Towler and Sinnott]]
'''Above Pinch'''
# CPh ≤ CPc
# We can match stream 1 or 2 with stream 3 but neither stream can match with stream 4. This creates a problem, since if we match stream 1 with 3 then stream 2 will not be able to make a match at the pinch. Likewise, if we match stream 2 with 3 then stream 1 will not be able to make a match at the pinch.
# Check the heat available in bringing the hot streams to the pinch temperature.
#::stream 1 ΔH = 40.0(180 − 100 )= 3200 kW
#::stream 2 ΔH = 30.0 (150 − 100 ) = 1500 kW
# Check the heat required to bring the cold streams from the pinch temperature to their target temperatures.
#::stream 3 ΔH = 60.0(180 – 80)= 6000 kW
#::stream 4 ΔH = 20.0(160 – 80) = 1600 kW
# If we split stream 3 into two branches with CP of 40.0 and 20.0, then we can match the larger branch with stream 1 and transfer 3200 kW, which satisfies (ticks off) stream 1.
# We now have two cold streams, both with CP of 20.0, and one hot stream (2) with CP of 30.0. We need to split stream 2 into two branches. As an initial guess these can both have CP of 15.0. We can then match one branch of stream 2 with the smaller branch of 4 and transfer 750 kW, and the other branch with stream 3, also for 750 kW, which then ticks off stream 2.
# Include a heater on the larger branch of stream 3 to bring it to its target temperature:
#::ΔHhot = 40 (100 ) − 3200 = 800 kW
# Include a heater on the smaller branch of stream 3 to provide the balance of the heat required:
#::ΔHhot = 20 (100 ) – 750 = 1250 kW
# Include a heater on stream 4 to provide the balance of the heat required:
#::Δ Hhot = 1600 − 750 = 850 kW
# Check sum of heater duties = 800 + 1250 + 850 = 2900 kW = hot utility target.
'''Below Pinch'''
# CPh ≥ CPc
# Note that stream 4 starts at the pinch temperature and so cannot provide any cooling below the pinch.
# We cannot match stream 1 or 2 with stream 3 at the pinch.
# Split stream 3 to reduce CP. An even split will allow both streams 1 and 2 to be matched with the split streams adjacent to the pinch, so try this initially.
# Check the heat available from bringing the hot streams from the pinch temperature to their target temperatures:
#::stream 1 ΔH = 40.0 (100 − 40 ) = 2400 kW
#::stream 2 ΔH = 30.0 (100 − 60 ) = 1200 kW
# Check the heat required to bring the cold streams from their source temperatures to the pinch temperature:
#::stream 3 ΔH = 60.0 (80 − 30 ) = 3000 kW
#::Stream 4 is at the pinch temperature
# Note that stream 1 cannot be brought to its target temperature of 40 °C by full interchange with stream 3 as the source temperature of stream 3 is 30 °C, so ΔTmin would be violated. So transfer 1800 kW to one leg of the split stream 3.
# Check temperature at exit of this exchanger:
#::Temp out= 100 – 1800/40= 55 °C, satisfactory
# Provide cooler on stream 1 to bring it to its target temperature; the cooling needed is
#::ΔHcold = 2400 − 1800 = 600 kW
# Transfer the full heat load from stream 2 to second leg of stream 3; this satisfies both streams.
==References==
Biegler LT, Grossmann IE, Westerberg AW. Systematic Methods of Chemical Process Design. Upper Saddle River: Prentice Hall; 1997.

Seider WD, Seader JD, Lewin DR. Process Design Principles: Synthesis, Analysis, and Evaluation. 3rd ed. New York: Wiley; 2004.

Towler G, Sinnott R. Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013.

Turton R, Bailie RC, Whiting WB, Shaewitz JA, Bhattacharyya D. Analysis, Synthesis, and Design of Chemical Processes. 4th ed. Upper Saddle River: Prentice-Hall; 2012.

File:Composite Curves example.png

2015-03-02T02:58:53Z

Mzuckerman: Mzuckerman uploaded a new version of "File:Composite Curves example.png"

Pinch analysis

2015-02-07T04:59:55Z

Mzuckerman:

File:Cascade.PNG

2015-02-07T04:35:16Z

Mzuckerman:

File:Composite Curves example.png

2015-02-07T03:07:49Z

Mzuckerman: