Supporting Information for. Sodium Hydroxide Production from Seawater Desalination Brine: Process Design and Energy efficiency

Transcription

1 Supporting Information for Sodium Hydroxide Production from Seawater Desalination Brine: Process Design and Energy efficiency Fengmin Du a, David E. M. Warsinger a, Tamanna I. Urmi a, Gregory P. Thiel a, Amit Kumar a, John H Lienhard V a* a Rohsenow Kendall Heat Transfer Laboratory, Department of Mechanical Engineering Massachusetts Institute of Technology, 77 Massachusetts Avenue, Cambridge MA USA * Corresponding Author lienhard@mit.edu The supporting information gives Aspen modeling details of the brine-to-caustic process. Firstly, modeling tools available in Aspen are briefly introduced. Then, detailed modeling procedures of the membrane electrolyzer are given, including the validation of the developed model against a set of reference plant data in the literature. Dependencies of the electrolyzer performance on its modeling parameters are shown in a set of parametric studies. Afterwards, the Aspen models of pre- and post-treatment components involved in the brine-to-caustic process are elaborated. At last, mass flows, temperatures and salt concentrations at each stage of the process are summarized, which should help to reproduce the study. Additionally, we discuss briefly how the thermodynamic least work is defined and how it is evaluated in Aspen. 24 Pages, 7 Tables, 6 Figures S-1

2 Short Introduction to Aspen Plus Aspen Plus provides various features for construction and study of models. Some of them involved in this work are briefly introduced in this section. Blocks The following blocks (components) provided by Aspen Plus are implemented in our modeling. Their functions used are briefly introduced in the following: Mixer: allows for two or more feed streams to be combined; FSplit (Splitter): allows the feed stream to be separated into two or more fractions with same composition and properties; Sep (Separator): allows the feed stream to be separated to multiple fractions based on user-specified splits for each individual species; Flash2 (Flash): allows vapor-liquid separation by specified temperature, pressure or heat duty; Heater: models one side of a heat exchanger by changing the thermodynamic condition (temperature, pressure, vapor fraction) of the feed stream. Heat stream can be specified as well; HeatX (Heat exchanger): models two-stream heat exchanger; MHeatX: models multiple-stream heat exchanger; RStoic (Stoichiometric Reactor): models reactor with known extent or conversion; RGibbs (Gibbs Reactor): models equilibrium reactor. Aspen calculates the product composition by minimizing the Gibbs free energy; Pump: models pump; Compr: models compressor; Valve: models valve, i.e. for throttling; CFuge: models the filtration. S-2

3 Streams Streams either connect Blocks together or is used as feed or product of the flowsheet. Material, heat or work streams can be specified in Aspen plus, whereas only material streams are involved in this work. The composition, flow rate and thermodynamic properties have to be provided for feed material streams. 1 Other streams are calculated by Aspen based on the Blocks that they connect to. Internal recycling streams (i.e. the recirculating brine in Figure 1) are known as tear streams in Aspen. Aspen automatically iterates to solve for flows and compositions of these streams. However, initial guess of the composition and properties of the tear streams may be given by user to enhance the numerical stability of the model. Tools for simulation and analysis Design Specification (DS) A DS sets a variable in a simulation which would be calculated elsewise. Aspen internally adjusts a user-specified input parameter to meet this requirement. 1 For example, the requirement is the ph after an acidifier (mixer), and the requirement can be met by varying the acid stream into the mixer. Sensitivity Analysis Sensitivity analysis varies an input parameter and calculates how other user-specified variables depend on this parameter. Modeling details Membrane electrolyzer in detail Cathode chamber and membrane transport The transport through ion-exchange membrane is closely related to the production on the cathode side (red part in Figure 2, main paper), as the sodium ions that pass through the membrane become part of the product caustic. Along with sodium ions, water is transported through the ion-exchange membrane. 2 The molar ratio of S-3

4 water/sodium transported is called the water transport number, denoted by ff. Transport of other ions are of very small extent 2 and thus neglected. On the cathode, hydrogen and hydroxide ions are produced. Due to lack of cathodic side reactions, hydrogen produced by the cell is stochiometric and obeys Faraday s law: nn H 2 = nn H 2,id 1 = 2nn H 2 (II/FF) (S-1) Hereby, II is the current and FF the Faraday s constant (96,485 C/mol). The ideal stoichiometric production of NaOH is: 1 = nn NaOH,id (II/FF) (S-2) However, a fraction of the produced hydroxide ions undesirably migrates back to the anolyte through the membrane. This hydroxide back-migration reduces sodium transport through the membrane and the production of caustic soda. The cathode current efficiency ηη, which is defined by the following, 2 takes this effect into account: ηη = nn NaOH = nn NaOH nn NaOH,id (II/FF) = nn Na + (II/FF) (S-3) nn NaOH is the caustic soda molar stream produced in the cell. This equals to the sodium molar flow transported through the membrane, nn Na +. Anode chamber The reactions occurring are illustrated in Figure S1. Ideally, a Faraday of electricity (1 mol electrons) produces 0.5 mol gaseous chlorine. In real cells, however, this amount is reduced by oxygen production (side reaction 1 in Figure S1) and dissolution of chlorine into the anolyte (side reaction 2 in Figure S1). We take this into account by introducing the anode current efficiencies. S-4

5 Side reaction 1 Side reaction 2 Figure S1. Reaction network occurring in the electrolyzer anode chamber. The main reaction is chlorine oxidation (leading to Cl 2 ). The primary side reactions are water oxidation (1) and chlorine dissolution (2) in the anolyte solution. The blue color indicates reactions that occur on the anode and are irreversible. The green color indicates equilibrium reactions in the anolyte solution. The process chlorine current efficiency, ξξ P, is defined by equation (S-4), derived from the literature: 2 ξξ P = nn Cl 2,total nn Cl 2,id = 2nn Cl 2,total (II/FF) = nn NaCl (II/FF) (S-4) The chlorine current efficiency, also known as the anode current efficiency 2, ξξ, is defined by equation (S-5): ξξ = nn Cl 2,g = 2nn Cl 2,g nn Cl 2,id (II/FF) (S-5) where nn Cl 2,total is the total chlorine flow produced on the anode while nn Cl 2,g is the gaseous chlorine flow. nn Cl 2,id denotes the chlorine production of an ideal cell. The quantity ξξ P therefore accounts for side reaction 1 (water oxidation) whereas ξξ considers both side reactions 1 and 2 (water oxidation and chlorine dissolving). The oxygen production can be calculated based on ξξ P : nn O 2 = 1 4 II FF 2nn Cl 2,total = ξξp II FF (S-6) S-5

6 The NaCl conversion, XX NaCl, is the fraction of feed NaCl that reacts in the electrolysis cell and is a primary performance parameter. Equation (S-7) shows the relationship between this quantity and the electricity consumed: XX NaCl = nn NaCl = 2nn Cl 2,total = ξξp II nn NaCl,F nn NaCl,F nn NaCl,F FF (S-7) where nn NaCl,F is the molar flow of NaCl in the electrolyzer feed, and nn NaCl is the molar amount of NaCl that reacts in the electrolyzer per time unit. The process chlorine efficiency ξξ P plays a key role in both equations (S-6) and (S-7), linking the NaCl conversion and the oxygen production to the current consumed. For this reason, ξξ P is used instead of ξξ (in contrast to the literature 2 as the anode current efficiency in the electrolyzer modeling. Modeling of membrane electrolyzer in Aspen Plus Several modeling approaches of membrane electrolyzer are found in the literature. 3-6 While some investigated global performance of the membrane cell, 4,5 other focused on details i.e. reaction kinetics and transport phenomena. 3,6 As only steady-state solution is of interest here, we use a different modeling approach for the electrolyzer, namely by combination of available unit operations (Blocks) in Aspen. Without the modeling of reaction kinetics and transport phenomena, this simulation approach still allows the implementation of important process parameters (i.e. current efficiencies, water transport number) and the determination of material flows inside the cell in detail, while taking chemical and phase equilibrium into consideration. The resulting Aspen flowsheet is illustrated in Figure S2. S-6

7 Figure S2. Membrane chlor-alkali electrolysis cell model in Aspen Plus. The used unit operations and their major tasks are listed in Table S1. Table S1. Aspen units and their major tasks in the membrane electrolysis cell. Aspen block Corresponding cell part Major task of the unit Mixer ACIDIF Acidifier Adjust the feed brine ph RStoic R-AN Anode Reaction Cl to Cl 2 and H 2 O to O 2 /H + RGibbs REX Anolyte Equilibrium gaseous Cl 2 with anolyte Sep MEMB Membrane Separate Na + and H 2 O from anolyte -> catholyte RStoic R-CA Cathode Reaction H 2 O to H 2 /OH Sep S-BACK Membrane Separate OH from catholyte -> anolyte Flash2 FLASH Catholyte Separate gaseous hydrogen from catholyte FSplit CAUS-SPL - Split concentrated caustic into product and recycle stream Mixer CAUS-MIX - Mix the recycle stream with deionized water S-7

8 Modeling parameters used are listed in Table 1 (main paper). The Design Specs (DS) used in the model are summarized below: Table S2. Aspen Design Specs in membrane electrolysis cell. Aspen Design Spec Expression Vary Anode efficiency Cathode efficiency nn Cl 2 nn H 2 = ξξ P = 96% nn H 2 nn Na + nn H 2 = ηη = 94% nn Na + Caustic concentration in ww NaOH,in = 30.3% Split ratio Caustic concentration out ww NaOH,out = 32% Dilution water flow O 2 production nn H 2 nn Cl 2 = 2nn O 2 nn O 2 Water transport number nn H 2 O nn Na + = ff = 4.25 nn H 2 O Hydroxide back-migration 2nn H 2 nn Na + = nn OH nn OH Validation of membrane electrolyzer model The results of the membrane electrolyzer model are validated against a set of reference plant data given in the Handbook of chlor-alkali technology. 2 Model parameters and their values for the validation are listed in Table S3. Since all model parameters except the mass flows in the feed are independent of mass, the model is scalable. For this comparison, all flows in the reference plant are scaled down by a factor of Table S3. Parameters used in model validation of the membrane electrolyzer. The quantity ξξ P affects the chlorine/oxygen ratio, ηη the hydroxide back-migration, and ff the water transport through the membrane. The fraction of feed NaCl that is converted in the cell is XX NaCl. Symbol Description Value Used Unit Source TT cell Temperature 88 C Reference 2 S-8

9 pp An Pressure anolyte 1.09 bar Reference 2 pp Ca Pressure catholyte 1.05 bar Reference 2 mm NaCl,F NaCl flow in the feed kg/h Reference 2 mm H 2 O,F Water flow in the feed kg/h Reference 2 mm HCl,F HCl flow in the feed kg/h Reference 2 mm Na 2 SO 4,F Na 2 SO 4 flow in the feed kg/h Reference 2 mm NaClO 3,F NaClO 3 flow in the feed kg/h Reference 2 ww NaOH,P Product NaOH concentration 32 wt% Reference 2 ww NaOH,F Recycle NaOH concentration 30.3 wt% Reference 2 ηη Cathode current efficiency 94 % Reference 2 ξξ P Anode current efficiency 95.5 % Fitted ff Water transport number Reference 2 XX NaCl NaCl conversion 44 % Fitted Table S4. Comparison of model with data from the reference plant 2 scaled down 1000 times. System characteristic Model Reference Unit Relative difference (%) Anolyte Flow kg/h +0.3 Anolyte Concentration wt% +0.1 Anolyte dissolved Cl 2 * kg/h Chlorine in anode gas kg/h 3.7 Oxygen in anode gas kg/h +1.8 Water in anode gas kg/h 15.3 Water flow through membrane kg/h +2.2 S-9

10 Hydrogen in cathode gas kg/h +0.9 Water in cathode gas kg/h 1.6 Caustic production kg/h 0.8 *Anolyte dissolved Cl 2 counts equivalent amounts of ClO (1 mol equivalent to 1 mol Cl 2 ) and ClO 3 (1 mol equivalent to 3 mol Cl 2 ). The results are shown in Table S4. Most of the relative differences are under ±5%. The electrolyzer model in Aspen is therefore capable of predicting the stream properties from and to the membrane cell under given parameters. Additionally, the validation confirms the scalability of the model. However, two items in Table S4 have high error: the Cl 2 dissolved in the anolyte and the water in the anodeside vapor. The latter does not affect the system performance and is thus tolerated. The former is presumably caused by the assumption that Cl 2 is in complete chemical equilibrium with the anolyte (side reaction 2 in Figure S1). In a real system, the ideal equilibria may not be reached due to finite reaction rates. Modeling of the reaction kinetics is beyond the detail level of this system-level modeling study. The influence of this error on the system process performance would be simply the overestimation of the dechlorination chemical dosage. Another possible reason for the deviation between simulation and real values is the inaccuracies in the thermodynamic property functions, implemented in the Aspen ENRTL model. Parametric study of the membrane electrolyzer model This section gives an overview of the dependency of membrane cell performance on its parameters, introduced in Table 1 (main paper) and SI section Modeling of membrane electrolyzer in Aspen Plus. All parametric studies of this section are based on the validation case, introduced in the last section. NaCl conversion in electrolyzer NaCl conversion is the fraction of NaCl reacted in the cell. Due to stoichiometry of the electrolysis reaction, the NaCl conversion directly relates to the current consumption in the cell (see equation (S-7)). In the parametric study, the NaCl conversion is varied from 0.4 to 0.5. Figure S3 shows product (NaOH and Cl 2 ) productivity at the anode compared to the NaCl conversion: the relationship is linear. S-10

11 Figure S3. Productivity of the electrolyzer (production rate of NaOH and gaseous Cl 2 ) with respect to NaCl conversion. The stability limit (dotted line) on anolyte concentration is 200 g/l NaCl (= 18.2 wt% at 90 C, by Aspen). 2 Other parameters remain as the validation case (see Table S3). Although cell productivity increases with NaCl conversion, in the given case, a >49% conversion would lead to unstable operation (red zone). Higher conversion (shown in green, right axis) values are desirable but are limited by instability at low anolyte NaCl concentration (unstable below 18.2 wt%), limiting conversion to no more than 49%. In the process chain, the NaCl conversion is adjusted by a DS so that a 19 wt% anolyte concentration is reached (see Table 1 in main paper). This way, a relative high cell productivity is achieved with a safe margin from unstable operating conditions. Anode and cathode current efficiencies in electrolyzer Current efficiencies play a key role in terms of energy consumption and cell productivity of the electrolyzer and require a detailed parametric study. Varying the cathode current efficiency ηη between 90% and 100% yields Figure S4. S-11

12 Figure S4. Impact of cathode current efficiency on NaOH and gaseous Cl 2 production as well as anolyte ph, anode current efficiency 95.5 % constant, other model parameters as Table S3. The ph declines rapidly at current efficiencies above 95%, risking system damage. Figure S4 shows the cell productivity as well as the ph of the anolyte solution with respect to the cathode current efficiency ηη. NaOH production increases linearly with an increasing efficiency. This can be deduced from equation (S-3). At low cathode current efficiencies (ηη < 95%), chlorine gas production is not at maximum. This is due to the back-migrated hydroxide ions which react with the chlorine. As the hydroxide ions are always fully depleted by chlorine, the ph value of the anolyte does not go beyond 4.3. At high cathode current efficiencies (ηη > 95%), however, the ph of the anolyte goes down to almost 1, which is undesirable in membrane cells as the ionexchange membrane may be protonated. 2 It is therefore crucial to monitor and control the ph at the anolyte by adjusting the HCl dosage in the brine acidification. Literature 2 suggests that, at 96% cathode current efficiency or higher, very little or no acid is needed. This agrees well with our result. S-12

13 The results of varying the anode current efficiency ξξ P are illustrated in Figure S5. While varying ξξ P, the total consumed electric current II is kept at constant. The NaCl conversion is adjusted accordingly, based on equation (S-7). Figure S5. Impact of anode current efficiency on production of NaOH and gaseous Cl 2 as well as the anolyte ph under constant current consumption, cathode efficiency 94% constant, other parameters as Table S3. It can be concluded that caustic soda production under constant current consumption is independent of ξξ P. Chlorine gas production as well as the anolyte ph reach a plateau at 94% anode current efficiency and do not increase further. The transition occurs exactly when both current efficiencies are equal: ηη = ξξ P (S-8) This transition is also present in Figure S4 where the anolyte ph and chlorine production largely change when ηη reaches about 95%. Note in that sensitivity study, the anode current efficiency is 95.5%. S-13

14 This fact can be explained by comparing the protons produced on the anode and the hydroxides leaked back to the anode. If anode (ξξ P ) and cathode efficiencies (ηη) are equal, the protons and hydroxides exactly neutralize. nn OH = (1 ηη) II FF ; nn H + = 4nn O 2 = 1 ξξ P II FF (S-9) If ηη is higher, protons will be in an excess amount which will bring the anolyte ph down. If ξξ P is higher, hydroxides will remain and react with chlorine, reducing the gas production amount. It can be concluded that high anode and the cathode current efficiency is beneficial for chlorine and caustic soda production, respectively. However, both efficiencies ought to have close values from each other. If ηη is too low, the chlorine produced will be wasted in the cell and the dechlorination costs will increase. If ξξ P is too low, the anolyte ph will be very low and hard to control. Given the results of the sensitivity study, the feed brine ph in the final system-level process is set at 3. Although lower feed ph brings improvement in the anode current efficiency 2, a ph under 2 is not tolerable for the cell. 2 As ph of the anolyte is strongly dependent on the relative value of both current efficiencies which are not controllable, a safe distance to the limit should be kept. The chosen brine ph is slightly lower than that in the validation case. A lower feed brine ph leads to a higher anode current efficiency 2 and an estimated value of 96% compared to 95.5% in the validation case is applied (see Table 1 of main paper). Varying feed NaCl concentration into the electrolyzer The feed NaCl concentration is an important parameter in the electrolyzer as it determines the cell productivity and the energy consumption from the concentration steps earlier. Its impact on the cell itself is investigated in this section. In the sensitivity study, feed NaCl concentration is varied by changing the water flow from Table S3. NaCl flow and other parameters are kept constant. The change in water flow matches a NaCl solution of 25 wt% to 27 wt% (slightly oversatured, fed with solid NaCl). The interval is chosen according to the feed concentration requirement of g/l NaCl. 7 Apparently, a higher feed NaCl concentration would result in a higher anolyte concentration, providing that the current consumed is kept constant. As the maximum NaCl conversion is majorly dependent on the lower limit of the anolyte NaCl concentration, a higher feed NaCl concentration makes higher conversion possible. The S-14

15 maximum cell productivity (as dry NaOH) at maximum conversion (anolyte at 200 g/l, limit of stable operation) is compared in Figure S6 where it is shown that the cell productivity is higher with feed NaCl concentration. Figure S6. NaOH production rate (dry) at various feed brine concentration, maximum conversion (anolyte concentration 200 g/l), model parameters as Table S3. Modeling of other components in Aspen Plus Summary of components Table S5 lists the individual pre- and post-treatment components with their main roles in the process chain. Table S5. Summary of components and their major tasks in the process chain (Figure 1). Component Category Major Task Nanofiltration (NF) Primary purification Remove some of Ca, Mg; remove most of sulfate S-15

16 Electrodialysis (ED) Primary concentration Concentrate up to 20 wt% NaCl Evaporation or Mechanical Vapor Compression (MVC) Final concentration Concentrate up to NaCl saturation Chemical Softening Secondary purification Remove Ca, Mg to ppm-level Ion-Exchange (IX) Final purification Remove Ca, Mg to ppb-level Dechlorinator Post-treatment Remove free chlorine produced in the electrolyzer Brine acidifier Brine acidifier is the pretreatment step immediately before the membrane electrolyzer. As ion-exchange works best at elevated ph, 2 the acidifying should be done only after the IX step. The acidification by adding HCl increases the anode current efficiency of the cell 2 and decomposes carbonate and hydroxide left in the chemical softening stage. The optimal ph after the acidification is The acidifier is modeled by two Mixers. HCl is dosed to the brine to decompose the carbonate and hydroxide left in the chemical softening stage in the first Mixer. The amount of acid is set to be stoichiometric. In the second Mixer (implemented in the electrolyzer model), the acid is used to adjust the brine ph. The amount of HCl is set by a Design Spec to obtain the desired ph. Ion-exchange Cation ion-exchange (IX) is the final purification step before the electrolyzer and is used to remove Ca 2+ /Mg 2+ as well as other possible cation impurities from the feed brine. The ion-exchange resin for brine purification must specifically be highly selective to Ca 2+ /Mg 2+ ions so that the desired Na + ions in the brine are not removed. In this study, the resins are modeled with the properties of the resin Amberlite TM IRC747 from Dow. 8 The basic reaction for IX columns is: 2 M R Na + (R ) 2 M Na + (S-R1) where M 2+ can be any divalent ions (Ca, Mg, Sr, Ba, etc.). The species R denotes the polymer bone of the resin with a negative charge. The resin needs to be regenerated back to full capacity after it becomes (near- )saturated with impurities. HCl and NaOH are consumed to regenerate the resin. 2,8 S-16

17 The IX column is modeled as a continuous process like other components in Aspen. Resin and regenerant (HCl, NaOH) amount needed are estimated based on technical datasheets from the supplier. 8 Following assumptions are made: Ion-exchange reactions have 100% conversion. This leads to a zero-hardness level in the electrolyzer feed. In the real system, there might be a ppb-level of hardness leakage. However, this difference does not affect modeling results in the electrolysis cell. The regeneration of resin is complete. Temperature of brine does not change in IX stage. Chemical softening The chemical softening stage includes dosage of soda ash (Na 2 CO 3 ) and caustic soda. Two major precipitation reactions are present for the removal Ca 2+ and Mg 2+ : Ca 2+ + CO 3 2 CaCO 3 (s) Mg OH Mg(OH) 2 (s) (S-R2) (S-R3) The model in Aspen uses two separate Mixers and a CFuge (Filter) in Aspen. Typically, a certain over-dosage of the chemicals is necessary to achieve a low remaining concentration of both Ca 2+ and Mg 2+. We apply the values given by reference 2, which are 0.6 g/l NaOH and 0.3 g/l Na 2 CO 3. In Aspen, mass flows of both precipitants are controlled by two design specs (DS), varying concentration of carbonate and hydroxide in the effluent of the chemical softeners. The precipitate is separated from the brine by a perfect filtration. Evaporation A steam-driven evaporator is one of the possibilities to concentrate the ED concentrate to saturation. The evaporator is modeled by using a Heater and a Flash2 in Aspen. In the Heater, the brine is isobarically heated to a vapor-liquid-mixture whereas in the Flash2, both phases are separated adiabatically and isobarically. A DS adjusts the vapor fraction at the outlet of the Heater to achieve the desired brine outlet concentration (set as 27 wt%). A preheater (as HeatX, two-stream heat exchanger) is implemented to preheat the ED concentrate before mixing it with the recycle stream. The hot side of the preheater is the hot brine at the outlet of the Evaporator. The output brine is cooled to 60 C, temperature for chemical softening. S-17

18 Mechanical vapor compression MVC is the alternative for the final concentration step. The MVC process is modeled by a Preheater (MHeatX), a boiler (HeatX and Flash2), a compressor (Compr), a Pump and a Valve. The feed stream is throttled in the Valve to a pressure of 0.3 bar, then preheated (to 65 C) and sent to the boiler. A DS (Design Spec) controls the vapor fraction after the boiler so that the desired brine concentration is achieved (27 wt%). The vapor-liquid-mixture is separated by the adiabatic Flash2 and the vapor gets compressed in the Compr. The compressed vapor is used to heat the boiler. The condensed vapor (hot water) and the hot brine are both put in the hot side of the preheater. The brine is cooled to a fixed 60 C for the chemical softening stage. A DS controls the compression ratio in the compressor to ensure that the hot water is cooled exactly to 60 C in the preheater. The cooled brine is brought back to ambient pressure by a Pump block and fed in the chemical softening stage. Electrodialysis The eletrodialysis process (ED) is modeled as a black-box model in Aspen. The model was an improved adaptation of a blackbox ED model developed by Nayar et al., 9 using the same equations for calculating energy consumption and costs as Nayar et al. 9 However, it is an improvement over Nayar et al. s blackbox model in that the effect of water transport across the ED membrane was also captured without the need for discretizing the stack in to smaller computational units. The Aspen model was further fine-tuned to approximate the results from a more sophisticated ED model developed by Nayar et al. 10 that simulated the performance across the length of an ED stack. The blackbox Aspen model assumed that there s only transfer of monovalent ions and pure water across the ED membrane. The effect of pure water transport was captured by fixing the ratio of mass flow rate between the concentrate outlet and inlet to be 4:1. This ratio was based on the simulation results from Nayar et al. 10 which accounted for water transport across the membrane and was in alignment with information we gathered from the industry. The energy consumption is calculated from the amount of salt transported by the current while accounting for a current utilization factor of 0.7 and a voltage of 0.3 V, similar to Nayar et al.. 9 WW ED = UU ED II ED = UU ED FF nn Na +,ED ηη ED (S-10) S-18

19 where UU ED = 0.3 V is the voltage of the ED stack, ηη ED = 0.7 the current utilization factor, which describes how much the total current effectively contributes to ion transport through ED membrane. nn Na +,ED is the molar amount of sodium ions transported through ED membrane. Nanofiltration The major role of nanofiltration is the removal of sulfate ions. In addition, NF removes a fraction of hardness ions (Ca 2+ and Mg 2+ ). However, NF also partially removes Na + and Cl ions. ROSA 9 from Dow 11 is used to model the NF system. The NF membrane 12 is chosen due to its low NaCl rejection and high active surface area. Based on the permeate flow needed for the ED process, the feed flow is calculated based on minimum retentate flow for the given membrane unit (provided by ROSA 9). The water recovery can then be determined as the ratio permeate/feed. A constant density is assumed during the calculation. ROSA 9 calculates the required feed pressure for the separation. Additionally, rejections of individual ions are obtained from the software, which are defined by: RR i = cc i,p cc i,f 100 % (S-11) Where RR i is the rejection of the ion i, and cc i,p and cc i,f refer to the concentration of the ion i in the permeate and the feed, respectively. The rejection of Br is not calculated by ROSA. Instead, it is assumed that Br behaves like Cl. The rejections as well as the water recovery ratio are implemented in Aspen as various DS, controlling a Sep block which represents the NF membrane. Dechlorination Dechlorination is used to remove free chlorine from the depleted brine produced by the membrane electrolyzer. Its first step includes adding HCl to shift the chlorine equilibrium (Figure S1) towards gaseous chlorine (As the chemical equilibrium shown in Figure S1 produces protons, addition of acids suppresses the chlorine dissolution while addition of bases enhances it.) and separating the chlorine gas, e.g., with air stripping. 2 The second step is adding reducing agents (i.e., sodium bisulfite NaHSO 3 ) to remove the remaining chlorine. The reaction can be written as: HSO 3 + Cl 2 + H 2 O 2Cl + 2H + + HSO 4 (S-R4) S-19

20 In the Aspen modeling, dechlorination units are simplified to two Mixers, one with HCl and the other with NaHSO 3. A local chemical reaction system is implemented including all the global reactions (salt dissociation, precipitation) and the new redox reaction (S-R4). Aspen calculates the chemical equilibrium based on all the reactions in the local system. The HCl amount is controlled by a DS setting the ph of the mixture. The NaHSO 3 amount is adjusted to the stoichiometric amount of chlorine after the primary dechlorination. No over-dose is needed here due to the complete redox reaction. Others Purge: The purge is implemented by a FSplit component which separates the stream to purge and recycled stream based on the user input split ratio. Brine heater: A Heater is used to heat the brine to the desired 88 C feed temperature of the membrane electrolyzer. Brine pump: A Pump is used to bring the brine from ambient (1 bar) to the anode side pressure in the membrane electrolyzer (1.09 bar). Summary of modeling parameters Table S6 lists modeling parameters not included in the main text. Table S6. Summary of detailed modeling parameters of the final system-level process not shown in the main text. Component Parameter Value Unit Source/Rationale Eletrodialysis Mass flux increase in C channel 4 - SI section Electrodialysis (C: concentrate, D: diluate) Chemical Softening Voltage 0.3 V SI section Electrodialysis Current efficiency 70 % SI section Electrodialysis NaOH concentration 250 g/l Reference 13 NaOH over-dosage 0.3 g/l Reference 2 Na 2 CO 3 concentration 37 g/l Reference 13 S-20

21 Na 2 CO 3 over-dosage 0.6 g/l Reference 2 HCl amount (as dry) 9.25 g/h Reference 8 + Calculation/Estimation IX HCl concentration 6 wt% Reference 8 NaOH amount (as dry) 8.22 g/h Reference 8 + Calculation/Estimation NaOH concentration 4 wt% Reference 8 Acidifier Dechlorination HCl concentration 37 wt% Commercial concentrated HCl HCl amount kg/h Acidify feed brine to a ph of 3 NaHSO 3 amount 4.32 kg/h Stoichiometry to free chlorine NaHSO 3 concentration 38 wt% Common strength 2 Summary of process outputs Detailed operating conditions (mass flow rate, temperature, salt mass fraction) at all stages of the process are given here, which will be helpful for reproducibility of this study. Table S7 summarizes the flows in each stage of the system-level process and some of their properties. Table S7. Summary flows of the final system-level process. Stream Mass flow Temperature Solution concentration [kg/h] [ C] [wt%] NF feed C 7.0 NF retentate C 7.5 NF permeate = ED feed C 5.7 ED diluate C 3.5 ED concentrate C 20.0 Recycle stream C 19.3 Evaporator/MVC feed C (after preheater) 19.9 Evaporator/MVC vapor C Pure water Evaporator/MVC brine = Chemical softening feed C 27.0 S-21

22 Chemical softening NaOH C 20.5 Chemical softening Na 2 CO C 3.6 Precipitate C Solid mixture Chemical softening out = IX feed C 26.2 IX HCl C 6 IX NaOH C 4 IX out = Acidifier feed C 26.2 Acidifier HCl C 37 Acidifier out =Electrolyzer feed C 26.2 Electrolyzer Cl C As pure gas Electrolyzer H C As pure gas Electrolyzer NaOH C 32 Electrolyzer depleted brine = Dechlorination feed C 19.0 Dechlorination NaHSO C 38 Purge C 19.3 The scalability of the model allows the application of other NaOH production rates. Least Work Energy use is a dominant cost and environmental concern for brine concentration and chlor-alkali electrolysis. For improving energy efficiency, the energy used must be compared to the thermodynamic least work, i.e., the minimum energy possible for a given process. The least work for an individual component in the system is given by: 14,15 S-22

23 WW least = Ξ out Ξ in = mm (h TT ss) mm (h TT ss) out in (S-12) where Ξ out and Ξ in are the exergy (available work) of the streams passing through the outlet and the inlet, TT refers to environmental temperature ( K), and h and ss are the mass-specific enthalpy and entropy of a stream than can be obtained from the Aspen model. References (1) Aspen Technology Inc, Aspen plus: Aspen plus user guide. Technical Report Aspen Technology Inc., (2) O Brien, T. F.; Bommaraju, T. V.; Hine, F. Handbook of Chlor-Alkali Technology. Springer, (3) Byrne, P.; Bosander, P.; Parhammar, O.; Fontes, E. A primary, secondary and pseudo-tertiary mathematical model of a chlor-alkali membrane cell. J. Appl. Electrochem. 2000, 30 (12), (4) McCluney, S.; Van Zee, J. An optimization analysis of a diaphragm cell/evaporator system for NaOH production. J. Electrochem. Soc. 1989, 136 (9), (5) Jalali, A.; Mohammadi, F.; Ashrafizadeh, S. Effects of process conditions on cell voltage, current efficiency and voltage balance of a chlor-alkali membrane cell. Desalination 2009, 237 (1-3), (6) Leah, R.; Brandon, N.; Vesovic, V.; Kelsall, G. Numerical modeling of the mass transport and chemistry of a simplified membrane-divided chlor-alkali reactor. J. Electrochem. Soc. 2000, 147 (11), (7) Bommaraju, T. V.; Lüke, B.; O Brien, T. F.; Blackburn, M. C. Chlorine. Kirk-Othmer encyclopedia of chemical technology, (8) Dow Water & Process Solutions, Amberlite TM selective resins for brine purification in the chlor-alkali industry, Technical Report Dow Water & Process Solutions. (9) Nayar, K. G.; Fernandes, J.; McGovern, R. K.; Dominguez, K. P.; Al-Anzi, B., Lienhard, J. H. Costs and energy needs of RO-ED hybrid systems for zero brine discharge seawater desalination. Int. Desalin. Assoc. World Congr. 2017, Sao Paulo, Brazil. S-23

24 (10) Nayar, K.; McGovern, R.; Fernandes, J.; Al-Anzi, B.; Lienhard, J. H. On the costs and energy benefits of hybridizing RO and ED systems for salt production. Desalination 2017, in Manuscript. (11) Dow Water & Process Solutions, Design Software: ROSA (Reverse Osmosis System Analysis). (revised on ). (12) DOW FILMTEC Membranes, Product data sheet of DOW FILMTEC TM NF /34i element. Technical Report DOW FILMTEC Membranes, (13) Garriga, S. C. Valorization of brines in the chlor-alkali industry. Integration of precipitation and membrane processes. PhD thesis, Universitat Politècnica de Catalunya, (14) Lienhard, J. H.; Mistry, K. H.; Sharqawy, M. H.; Thiel, G. P. Thermodynamics, Exergy, and Energy Efficiency in Desalination Systems. In Desalination Sustainability: A Technical, Socioeconomic, and Environmental Approach, Chpt. 4, H. A. Arafat, ed. Elsevier Publishing Co., (15) Mistry, K. H.; McGovern, R. K.; Thiel, G. P.; Summers, E. K.; Zubair, S. M.; Lienhard, J. H. Entropy generation analysis of desalination technologies. Entropy 2011, 13 (10), S-24