Desalination - Team A

Author: Pear Dhiantravan^[2016], Reed Kolbe^[2016], Sheridan Lichtor^[2016], John Marsiglio^[2016], Ellen Zhuang^[2016]

Instructors: Fengqi You, David Wegerer

Winter 2016

Introduction

An increasing global water scarcity is fueling initiatives everywhere for clean water treatment, making efficient seawater desalination an attractive aim for chemical plant design. A 2015 market analysis found that the global desalination market earned revenues of $11.66 billion, and this number is expected to reach over $19 billion by 2019. Furthermore, the 17,000 desalination plants currently in operation are expected to double in number by 2020.¹

As fresh water sources become increasingly scarce, strict water conservation measures are being observed. California is entering the fourth year of one of its most severe droughts on record. Cities are required to reduce their water usage by 35% to avoid facing fines.² In lieu of these pressing conditions, California is looking for ways to provide more accessible fresh water to its citizens. This high demand for clean water motivated our choosing Richmond, CA as our plant location. At Richmond, the desalination plant can convert readily available seawater from the San Francisco Bay into fresh water. We propose to desalinate 50 million gallons of water per day to be sent to water treatment plants for potability. This aim is based on the capacity of the desalination plant in Carlsbad, California, which services a similar population.

The salinity of the water in the San Francisco bay is seasonal. During the dry seasons of summer and fall, salinity is high around 10 PSU (10,000 ppm) because water from the Pacific Ocean flows into the San Francisco Bay. During the wet winter season, fresh water from the rivers flows into the Bay, and salinity drops to around 2 PSU (2,000).³ Designed to meet the “worst case” scenario, our plant will process 398 million gallons of water per day with 10,000 ppm of salt to produce 50 million gallons of water with 1,000 ppm of salt. This gives a yield of 13%.

The desalination plant proposed consists of 18 flash stages in series. This report outlines our investigation of this potential water desalination plant, including design and optimization of the desalination process, an analysis of its energy consumption and environmental impacts, and a study of its economic implications.

Technical Approach

For the desalination step of our process, we selected a multistage flash (MSF) distillation process. Not only is MSF distillation a very viable method, but it is also the most common method, currently producing about 60% of the world’s desalinated water. MSF and reverse osmosis (RO) are the two major methods being used in large-scale desalination plants. Both processes require considerable amounts of energy. RO typically has a lower energy demand; however, the high impurity content of the Bay water would frequently necessitate membrane cleaning and/or exchange. Feed going through a RO system requires extensive pretreatment to remove biological organisms and other solids to control the pH and the chemical composition of the water. With the San Francisco Bay as our source of water, MSF is an attractive option because sediment and other large impurities can be separated from the feed before distillation occurs. Additionally, MSF distillation plants can be located near power plants and paired to their waste heat streams to conserve energy. This can reduce energy needs by half or two thirds, making MSF increasingly more practical. Due to the maintenance required in the RO system, the MSF distillation is the better option in terms of operation cost.⁴

A HYSYS model was built to simulate the MSF distillation section of the process, which spans after seawater pretreatment and before local potable water treatment. The model is used to simulate the distillation steps and to calculate necessary heating, cooling, and input flow rates to produce the required 50 million gallons of distillated water per day. For the simulation, the 398 million gallons per day of seawater feed stream is defined to be composed of 3.31 mol% sodium chloride (NaCl) and 96.69 mol% water. The percentage of NaCl is set to account for other ions and chemicals present after the pretreatment phase. NaCl alone is added to the HYSYS system due to the capabilities of the fluid package used to simulate this process, ElectroNRTL. The feed is set at 22°C, which is the approximate the temperature of the San Francisco Bay year round.

Process Design

Design Overview

The proposed desalination plant uses MSF distillation to convert the San Francisco Bay water to fresh water. The final process design contains 18 flash stages in series. Overall, 398 million gallons per day of of brackish water is converted to 50 million gallons per day of deionized water, a 13% yield. The production meets the required production rate of 50 million gallons of freshwater per day.

Seawater is taken in from the San Francisco Bay and mixed with the brine recycle stream produced downstream. The feed is then heated through a series of condensers and a heater. Distilled water vapor in each flash stage contacts the tubes carrying the seawater feed and condenses into a liquid stream by exchanging heat with the cooler seawater feed stream. The seawater is sent through a heater to increase its thermal energy before entering the first stage of the MSF. In each flashing stage, some of the water evaporates, leaving the salt behind. The freshwater vapors condense, as described above, and is collected and sent to a water treatment plant. The brine stream is split into the recycle stream and a waste stream, and the waste stream is diluted with fresh seawater before being discharged back into the bay.

A process flow diagram of the process is depicted in Appendix A.

Pretreatment

Sediment is removed to prevent solids from plugging the process downstream. Scaling of the process units leads to decreased efficiency and more frequent unplanned downtime, and is controlled for by lowering pH. Acidifying the water will also remove CO2, which is a corrosive gas. To control for oxygen, oxygen scavengers will be included in the system to sequester it. Sodium bisulfate will be used as the scavenger.⁵

Condensers (Heat Exchanger Networks)

In order to conserve and recycle energy, the condensation of the distillate and heating of the seawater feed will be coupled in countercurrent heat exchanger networks. These heat exchanger networks are condensers of the design. The cold feed, which is a mixture of fresh seawater and brine recycle, is fed through multiple heat exchangers in series. Each heat exchanger corresponds to a different evaporation chamber. This process consists of 18 heat exchangers: E-101, E-102, … E-117, and E-118, which correspond to evaporation chambers V-102, V-103, … V-118, and V-119, respectively.

The heat exchangers can be viewed similarly to a shell and tube setup, however there is no actual physical shell. The tubes simply run through the top of the evaporation chambers, where vaporized distilled water in the chamber acts as the shell side fluid. The tube bundles are arranged in long tube configuration and are aligned parallel to the direction of the flashing brine flow in the evaporation chamber. The evaporating brine in an evaporation chamber rise up to the the corresponding heat exchanger and condense on the outside of the tubes, creating collectable pure water. This condensing water is what heats up the seawater in the tubes.

E-101 was used as the basis for the condenser sizing since it has the highest heat requirement, and thus will be the largest heat exchanger. From the HYSYS model, the overall system requires 2.24 x 10⁶ kW of duty. Assuming this energy requirement, 5000 tubes, and 18 m length tubes (the diameter of the evaporation chambers, explained further on the next page), the diameter of each tube was calculated to be 0.4 m. Based on the temperature of the streams and flow, the thickness of the tube is 7 x 10^-4 m.

Brine Heater

After the seawater exits the final heat exchanger, it enters a brine heater (E-119) where it is brought up to the temperature necessary for the flashing process to begin. This means that the seawater must be overheated compared to stage one, so that the seawater will flash in the evaporation chamber, i.e. release heat and vapor to reach equilibrium with stage conditions. After flowing through the condensers, the seawater enters the brine heater at 96.1 ^oC and is heated to 120 ^oC. The overall system requires 2.31 x 10⁶ kW of duty. This is supplied by a total of 4.15 x 10⁶ kg per hour of steam.

Evaporation Chambers

When the seawater has been brought to a temperature appropriate for flashing, it enters the first evaporation chamber (V-102). Upon entering the chamber, the water is flashed. Some of the water evaporates, leaving the salt behind. The vapor rises up to the heat exchanger tubes in each evaporation chamber and condenses on the outside into the distilled water product. This condensing water is what heats up the seawater in the tubes discussed above. The seawater that does not evaporate moves to the next stage, which operates at a higher temperature and lower pressure, and the process is repeated. This continues in series through a total of 18 flashing chambers (V-102, V-103, … V-118, and V-119). All of the distillate is collected into one stream and all of the brine is mixed with seawater before being disposed of back in the bay in order to dilute it enough so it does not have an adverse effect on marine life.

Each respective heat exchanger fully condenses the vapor product stream from each stage. However, when all of the individual product streams are consolidated into the final combined product stream, some water exists in the vapor phase due to temperature and pressure changes associated with mixing all of the streams together. Although the vapor product stream from each stage condenses completely across its respective heat exchanger, the total product stream contains some water in the vapor phase due to the temperature change when the mixer sums all the streams. In order to fully condense this stream, which will be sent to local water treatment plans, a heat exchanger is put in place to cool down the product using the process feed stream.

The correlations of El-Dessouky and Ettouney were employed in order to size the flash chambers. They specified that conventional MSF plants that operate at a capacity between 27,000 and 36,000 m³/day use flashing stages of the dimensions 18 x 4 meters.⁶

Valves

The model requires valves in between the evaporation chambers. The seawater streams flow through valves with large pressure drops to flash a portion the liquid into vapor that can then be separated in the vessels. Large pressure drops provide the thermodynamic driving force to vaporize more of the liquid, and are required to achieve high product yield.

Material of Construction

Before 1980, the shells and internals of the MSF units were commonly constructed from carbon steel. Seawater, however, corrodes carbon steel. To compensate for this, the thickness of the carbon steel has an additional corrosion allowance, which increases the size and weight of the equipment. After 1980, the stainless steel and duplex stainless steel became increasingly common. Stainless steel metal requires less thickness for equivalent strength and corrosion resistance, which allows for a reduction in the size, weight, and therefore cost of units. Major units in our process will be constructed from stainless steel.⁶

Process Alternatives

In order to ensure that our design is fully optimized, we wanted to consider several alternatives for our process. We looked at alternatives associated with water pretreatment, the distillation process itself and product treatment.

Water Pretreatment

Scaling of the process units leads to decreased efficiency and more frequent unplanned downtime. Scaling can be controlled by adjusting pH, temperature, and bicarbonate concentration. We selected acidification to control for process scaling. This is because lowering a process temperature would decrease overall efficiency, and decreasing bicarbonate concentration would be comparatively expensive. Lowering pH would have the added benefit of controlling CO2 concentrations since CO2 is a corrosive gas. To control for corrosion we looked at how to remove oxygen. To target the oxygen we chose an oxygen scavenger. This was chosen over other options such as a deaerator, or a protective coating because a deaerator would have a large capital cost, and the equipment itself would experience corrosion, and a tank coating such as zinc orthophosphate has a mixed effectiveness.⁷