Sulfuric acid production with mineral carbon sequestration

Abstract

A geomimetic process of sulfate replacement by mineralized carbonate, either in situ or ex situ, is used for mineral carbon sequestration and critical element recovery.

Claims

1. A system that couples sulfuric acid production to mineral carbon sequestration, the system comprising: an electrolyzer stack of one or more electrochemical cells comprising: an anode within an anode chamber containing an anolyte, a cathode within a cathode chamber containing a catholyte, and an anion exchange membrane separating the anode and cathode chambers; a mineralized carbonate production reactor configured to receive a hydroxide solution from the cathode chamber, to generate mineralized carbonate from a sulfate feedstock and CO.sub.2, and to return some or all of the reactor solution to the cathode chamber; and a sulfuric acid recovery module configured to receive sulfuric acid from the anode chamber.

2. The system of claim 1, wherein the system is configured as a continuous flow system.

3. The system of claim 1, wherein the mineralized carbonate production reactor is operably connected to a source of sulfate.

4. The system of claim 3, wherein the source of sulfate comprises solid calcium sulfate.

5. The system of claim 1, where the mineralized carbonate production reactor is operably connected to a source of CO.sub.2.

6. The system of claim 5, wherein the source of CO.sub.2 comprises air.

7. The system of claim 1, wherein the mineralized carbonate production reaction is configured to convert a gypsum to calcium carbonate according to the reaction:
CaSO.sub.4.Math.2H.sub.2O(gypsum)+2OH.sup.?+CO.sub.2(g).fwdarw.CaCO.sub.3(s)+SO.sub.4.sup.2?(aq)+3H.sub.2O(l).

8. The system of claim 1, wherein the anion exchange member is configured so that sulfate anion crosses the anion exchange membrane to the anode chamber where sulfuric acid is generated.

9. The system of claim 8, wherein the system is configured to maintain a relatively low concentration of base (OH.sup.?) in the catholyte relative to the concentration of acid (H.sup.+) in the anolyte by recirculating fluid from the mineralized carbonate production reactor through the cathode chamber.

10. The system of claim 9, wherein the system is configured to generate an acid concentration in the anolyte that is higher than the base concentration in the catholyte even though protons and hydroxides are produced at the same rate in the electrochemical cell.

11. The system of claim 1, wherein on the anode side of the system, aqueous sulfuric acid is recirculated at a constant rate through the anode chamber to allow for accumulation of sulfuric acid.

12. The system of claim 1, wherein the system is configured for hydrometallurgical extraction or recovery using sulfuric acid obtained from the sulfuric acid recovery module.

13. The system of claim 12, wherein the hydrometallurgical extraction or recovery comprises sulfuric acid leaching of lithium claystone or other magnesium silicate.

14. The system of claim 13, wherein the system is further configured to return the leachate post-lithium extraction to the mineralized carbonate production reactor to recycle the sulfate and produce mineralized carbonate therefrom.

15. The system of claim 1, wherein the system is configured for phosphoric acid production with mineral carbon sequestration.

16. The system of claim 15, wherein the system is configured for generation of phosphoric acid from rock phosphorus with calcium carbonate as the solid product as described by the reaction:
Ca.sub.5F(PO.sub.4).sub.3(fluorapatite)+5CO.sub.2(g)+5H.sub.2O(l).fwdarw.5CaCO.sub.3(calcite or aragonite)+3H.sub.3PO.sub.4+HF.

17. The system of claim 1, wherein the system is configured to include cyclic steps of electrochemical production of sulfuric acid at the anode and calcium hydroxide aqueous solution at the cathode, wherein the hydroxide solution is reacted with carbon dioxide to produce solid calcium carbonate.

18. The system of claim 17, wherein the system is configured to include cyclic steps of electrochemical production of sulfuric acid at the anode and calcium hydroxide aqueous solution at the cathode, wherein the hydroxide solution is reacted with carbon dioxide and calcium ion to produce solid calcium carbonate, wherein the sulfuric acid anolyte is recovered, concentrated as necessary to >70% H.sub.2SO.sub.4 and reacted with rock phosphorus to produce phosphoric acid, calcium sulfate, and HF, wherein the product calcium sulfate is returned to the process to produce calcium carbonate and sulfate solution, wherein the sulfate solution is returned to the electrochemical cell along with water to continue the cycle.

19. The system of claim 1, further configured to sequester carbon dioxide as calcium carbonate and produce sulfuric acid by reacting calcium sulfate solids with electrochemically produced hydroxide solution contacted with carbon dioxide directly from air or from a more concentrated source.

20. The system of claim 1, further configured to include one or more of a sulfuric acid concentration step, a step to recover hydrogen or energy from produced hydrogen using a fuel cell, a phosphoric acid production step, and valuable co-product recovery steps.

Description

BRIEF DESCRIPTION OF THE DRAWINGS

[0096] FIG. 1 is a schematic diagram of the steps of a process for mineral carbon sequestration by carbonation of gypsum with production of sulfuric acid.

[0097] FIGS. 2A-C depict the time evolution of pH at each process sampling point shown in FIG. 1 for experiments run at three different currents respectively: 0.01 A, 0.02 A, and 0.05 A.

[0098] FIG. 3 shows the rate of carbonation as a function of the rate of OH.sup.? addition to the mixed flow reactor for the experiment run at 0.01 A current. The strong linear correlation between the calculated carbonation rate (R.sub.carb) and the hydroxide addition rate is indicative of a condition that is rate-limited by base production at the cathode.

[0099] FIGS. 4A-B illustrate the process efficiency as a function of current in terms of the acid production rate in moles H.sub.2SO.sub.4 produced per minute (A) and the power consumption per mole of H.sub.2SO.sub.4 produced (B) in solutions of sodium, calcium, and magnesium sulfate and mixtures thereof.

[0100] FIGS. 5A-B shows the results of the batch mode electrochemical efficiency test (A) illustrating high Faradaic efficiency of the electrochemical cell in 1M Na.sub.2SO.sub.4 solution. Experimental results (B) showing the evolution of Faradaic efficiency for the process illustrated in FIG. 1 as a function of the ratio of measured aqueous SO.sub.4 concentration to aqueous OH.sup.? concentration, demonstrating that maintaining low pH (low concentration of OH.sup.?) in the cathode chamber maximizes the process efficiency.

[0101] FIG. 6 is a flow chart of the steps of an embodiment of the method for mineral carbon sequestration by carbonation of gypsum with production of sulfuric acid, use of the sulfuric acid for leaching of lithium claystone for lithium recovery, return of the lithium-free leachate for recycling of both water and sulfate.

[0102] FIG. 7 is a flow chart of the steps of a method for phosphoric acid production with mineral carbon sequestration using sulfate autocatalytically to produce intermediate gypsum, which is subsequently reacted to form mineral calcium carbonate.

[0103] FIG. 8 is a flow chart of the steps of a method for mineral carbon sequestration by carbonation of gypsum with production of sulfuric acid.

[0104] FIG. 9 is a flow chart of the steps of a method for phosphoric acid production with mineral carbon sequestration using sulfate autocatalytically to produce intermediate gypsum, which is subsequently reacted to form mineral calcium carbonate.

[0105] FIG. 10 is a flow chart of the steps of an embodiment of the method for mineral carbon sequestration by carbonation of gypsum with production of sulfuric acid, with additional production of valuable co-products that are incompatible with calcium carbonate and therefore released into the sulfate solution during calcium sulfate conversion to calcium carbonate.

[0106] FIG. 11 is a flow chart of the steps of an embodiment of the method for mineral carbon sequestration by carbonation of gypsum with production of sulfuric acid, with additional production of valuable co-products that are compatible with calcium carbonate. Additional process steps implement a sodium hydroxide leach of the calcium sulfate with co-product recovery prior to conversion of the gypsum to calcium carbonate.

DESCRIPTION OF PARTICULAR EMBODIMENTS OF THE INVENTION

[0107] Unless contraindicated or noted otherwise, in these descriptions and throughout this specification, the terms a and an mean one or more, the term or means and/or. It is understood that the examples and embodiments described herein are for illustrative purposes only and that various modifications or changes in light thereof will be suggested to persons skilled in the art and are to be included within the spirit and purview of this application and scope of the appended claims. All publications, patents, and patent applications cited herein, including citations therein, are hereby incorporated by reference in their entirety for all purposes.

[0108] Aspects of the invention include a process, such as a continuous flow process, for sulfuric acid production that permanently sequesters CO.sub.2, such as atmospheric air-derived CO.sub.2, as mineral carbonate, e.g., through the replacement of calcium or magnesium sulfate by calcium or magnesium carbonate. FIG. 1 is a flow chart illustrating an embodiment of the invention wherein solid calcium sulfate is introduced to a mineral precipitation reactor where it is converted to calcium carbonate by reaction with carbon dioxide from air and alkalinity produced in a two-chamber water electrolyzer. Effluent from the precipitation reactor is recirculated through the cathode chamber of the water electrolyzer, where sulfate liberated in reaction 1 crosses an anion exchange membrane to gradually accumulate sulfuric acid in a recirculating anolyte solution. During operation of the calcium-containing system, sulfuric acid and calcium carbonate are produced by reacting a calcium sulfate source with electrochemically produced hydroxide contacted with carbon dioxide derived from atmospheric air, although more concentrated sources of carbon dioxide can also be used. While the illustrated embodiment is shown in the context of gypsum and the production of calcite, the invention is not so limited, as other sources of sulfate may be employed to generate other types of mineralized carbonate, e.g., as described above.

[0109] An embodiment of the invention integrates the process illustrated in FIG. 1 into hydrometallurgical extractions of magnesium silicate minerals with sulfuric acid to produce magnesium carbonate or hydroxycarbonate solids, green hydrogen, and recycled sulfuric acid. An example of this embodiment is illustrated as a simplified flow diagram in FIG. 6 for lithium extraction from claystone. Step 10 illustrates sulfuric acid production, recirculation, and accumulation in the anode chamber along with base production in the cathode chamber of an electrochemical cell or stack of electrochemical cells. Alkaline solutions produced in the cathode chamber are flowed to the mineral production reactor in step 20 where magnesium sulfate is reacted with CO.sub.2 from air or another concentrated source and alkalinity to produce solid magnesium carbonate products. Sulfuric acid is recovered and optionally concentrated in step 30, and produced green hydrogen gas is recovered and optionally either concentrated or used in a fuel cell to generate electricity in step 40. Magnesium sulfate solution is introduced to the mineral production reactor from the extractive process shown in Step 50. The extractive process in various embodiments can involve extractions of critical elements such as lithium from magnesium silicate materials using sulfuric acid of different concentrations ranging from about 1 wt. % to 70 or more wt. %. Acidity is neutralized in step 50 producing a sulfate waste stream that is introduced to the mineral precipitation reactor in step 20, which allows for recycling of the sulfuric acid used in the extractive process.

[0110] In other embodiments, the method can be integrated into phosphate fertilizer production to generate calcium carbonate byproducts instead of calcium sulfate, substantially mitigating an important environmental impact of agricultural fertilizer production, illustrated in FIG. 7. In this embodiment, sulfuric acid and alkaline solution are produced by water electrolysis in step 60 in an electrochemical cell or stack of electrochemical cells. Sulfuric acid is produced and optionally recirculated in the anode chamber along, and a hydroxide solution is produced from a sulfate feed solution in the cathode chamber, as shown in FIG. 1. Alkaline solutions produced in the cathode chamber are flowed to the mineral production reactor in step 70 where calcium sulfate is reacted with CO.sub.2 from air or another concentrated source and alkalinity to produce solid calcium carbonate products. Sulfuric acid is recovered and optionally concentrated in step 80, and produced green hydrogen gas is recovered and optionally either concentrated or used in a fuel cell to generate electricity in step 90. Calcium sulfate is introduced to the mineral production reactor from phosphoric acid production process shown in Step 100. The production of phosphoric acid is accomplished by reacting produced sulfuric acid with phosphate rock, which produces solid calcium sulfate (as phosphogypsum). The waste phosphogypsum produced in step 100 is introduced to the mineral precipitation reactor in step 70, which allows for recycling of the sulfuric acid in phosphoric acid production and importantly avoids the accumulation of phosphogypsum waste.

[0111] In an embodiment illustrated in FIG. 8, carbon dioxide is introduced to the process prior to the mixed flow reactor by a separate gas contactor apparatus in step 120, in a process otherwise identical to the illustration in FIG. 1. There is a large energy cost associated with bubbling gas through water as required when the gas contacting and precipitation reaction steps are combined into one step (e.g. steps 20 and 70). For example, for a gas disseminator producing 3 ?m air bubbles, the pressure drop is estimated to be 13 PSI based on the Young-Laplace equation. An air contactor operated at a very small pressure drop reduces the energy required to remove carbon dioxide from low concentration sources such as air. The addition of a separate gas contactor step as illustrated in FIG. 8 applies to all other embodiments of the invention, and an example of the overall process for phosphoric acid production using a separate gas contactor apparatus is illustrated in FIG. 9.

[0112] In certain embodiments, the method illustrated in FIGS. 7 and 9 can further include a step for separation of valuable elements released into solution during the conversion of calcium sulfate to calcium carbonate as shown in FIG. 10. In step 240, the conversion of phosphogypsum to calcium carbonate releases elements that are compatible in sulfate minerals but incompatible in carbonate minerals. These co-products may be separated from the sulfate solution stream as in step 250 prior to reintroduction of the sulfate solution to the electrochemical cell or stack of cells.

[0113] To selectively recover elements that are compatible with the carbonate mineral precipitates, elements can be recovered prior to the mineral production reactor as illustrated in FIG. 11. Step 290 consists of a basic leach of the material, followed by step 300 to recover the carbonate-compatible elements of interest. The extracted leachate can then be introduced to the contactor in step 310 of the embodiment, followed by sulfate carbonation in step 320.

[0114] In other embodiments, multiple embodiments can be combined to recover valuable elements before and after calcium sulfate carbonation, with or without phosphoric acid production.

Examples

[0115] In an aspect we describe a continuous flow process for sulfuric acid production that permanently sequesters atmospheric air-derived CO.sub.2 as mineral carbonate through the replacement of calcium or magnesium sulfate by calcium or magnesium carbonate. FIG. 1 is a flow chart illustrating an embodiment of the invention wherein solid calcium sulfate is introduced to a mineral precipitation reactor where it is converted to calcium carbonate by reaction with carbon dioxide from air and alkalinity produced in a two-chamber water electrolyzer. Effluent from the precipitation reactor is recirculated through the cathode chamber of the water electrolyzer, where sulfate liberated in reaction 1 crosses an anion exchange membrane to gradually accumulate sulfuric acid in a recirculating anolyte solution. During operation of the calcium-containing system, sulfuric acid and calcium carbonate are produced by reacting a calcium sulfate source with electrochemically produced hydroxide contacted with carbon dioxide derived from atmospheric air, although more concentrated sources of carbon dioxide can also be used. Experiments were performed over a range of current densities (6.3-31.3 A/m.sup.2) to investigate process efficiency under different kinetic regimes, and the rate of acid and base production is shown in FIG. 2A-C. At higher currents, the supply of CO.sub.2 was rate-limiting, causing the catholyte solution to evolve towards high pH values, while at the lowest current, the production of base at the cathode limits the rate of mineral carbonation, and the catholyte chamber evolves towards a low steady state pH of 8-9, allowing for controlled calcium carbonate precipitation as aragonite. The carbonation rates observed in the lowest current experiment are shown in FIG. 3, which shows that the carbonation rate is approximately half the rate of base production at the cathode, because two equivalents of alkalinity are consumed per mole of CO.sub.2 mineralized to CaCO.sub.3. The rates and energy intensities of acid production are summarized in FIG. 4A-B for the process embodiment detailed in FIG. 1 for a series of different catholyte feedstock compositions including calcium sulfate, magnesium sulfate, and sodium sulfate and mixtures thereof in different concentrations. FIG. 5A shows that the Faradaic efficiency for the embodiment illustrated in FIG. 1 is high and the resulting energy intensity of acid production was low (0.2 kWh/mol H.sub.2SO.sub.4 in this experiment), demonstrating the high efficiency of this process compared to other industrial electrochemical processes. FIG. 5B explains the efficiency of this process by showing that maintaining a low pH in the catholyte significantly enhances process efficiency relative to typical electrochemical acid/base production by reducing OH.sup.? migration across the anion exchange membrane relative to SO.sub.4.sup.?. The invention enables efficient sulfuric acid production in a simple, two-chamber electrochemical cell (<0.4 kWh/mol H.sub.2SO.sub.4, with greater than about 80% Faradaic efficiency).

[0116] Materials and Methods

[0117] Reactor Design and Operation

[0118] A continuous flow reactor system was developed that couples sulfuric acid production to mineral carbon sequestration using gypsum as the source of calcium and sulfate. A flow diagram of the system constructed for the tests described here is illustrated in FIG. 1. Briefly, a membrane separated electrochemical cell is connected to a mixed flow reactor, where gypsum is converted to calcium carbonate (reaction 2). Effluent from the mixed flow reactor is recirculated into the cathode chamber to supply sulfate anion, which crosses the anion exchange membrane to the anode chamber where sulfuric acid is generated. The hydroxide solution produced in the cathode chamber is returned to the mixed flow reactor, where it reacts with CO.sub.2 that is introduced by bubbling atmospheric air (reaction 3).

[0119] The electrochemical cell used in these tests consists of a platinized titanium anode and cathode separated by FuMA-Tech Fumasep FAS-PET-130 anion exchange membrane (AEM), connected to a potentiostat operated at a constant current, which are all readily available materials. The eluent of the mixed flow reactor is passed through a 0.45 um filter to remove suspended solids and then pumped into the cathode chamber of the electrochemical cell, where alkalinity increases by the hydrolysis reaction on the cathode: H.sub.2O+e.sup.?.fwdarw.OH.sup.?+? H.sub.2(g). The more basic effluent from the cathode chamber is returned into the mixed flow reactor to supply alkalinity for calcium carbonate precipitation via reaction 1, which lowers the solution pH as alkalinity is consumed. On the anode side of the system, aqueous sulfuric acid was recirculated at a constant rate through the anode chamber to allow for the gradual accumulation of sulfuric acid using pure water as a starting solution, where protons and oxygen are generated by hydrolysis at the anode. Sulfate ions migrate from the catholyte via the AEM to balance charge, and sulfuric acid is recovered as the anolyte.

[0120] To initiate the process, the cathode chamber (50 mL) and mixed flow reactor (50 mL) were first filled with aqueous solution pre-equilibrated with gypsum and atmospheric CO.sub.2, and then 3.0 g of powdered calcium sulfate dihydrate (gypsum) was added to the mixed flow reactor. The gypsum powder was prepared by crushing and milling selenite gypsum and sifting to recovery the <180 um size fraction (Ward's scientific). The initial mass of gypsum used was chosen such that the process rates are independent of the gypsum mass, as gypsum rapidly obtains chemical equilibrium with the aqueous solution. The pre-equilibrated aqueous solution was prepared by mixing doubly deionized water with 5.0 g of powdered gypsum and bubbling the solution with air for 30 minutes. The equilibrated solution was then vacuum filtered through at 0.2 ?m PTFE membrane (Millipore) before filling the mixed flow reactor and the cathode chamber. The anode chamber (also 50 mL) is initially filled with doubly-deionized water and connected to a recirculating reservoir with an additional volume of 50 mL, such that the anode and cathode sides of the system both have a total volume of 100 mL of aqueous solution.

[0121] To begin an experiment, the potentiostat was powered on at the selected current, and flow was initiated on the cathode and anode sides of the system using two pumps operated at different flow rates: anolyte solution was rapidly recirculated to reduce charge polarization, while the catholyte flow rate of ?3-5 mL/min was set to allow for an approximately 10-15 minute fluid residence time in the mixed flow reactor. The mixed flow reactor was sparged continuously with atmospheric air using a stainless-steel disseminator, which creates small bubbles that facilitate CO.sub.2 dissolution into the aqueous solution. The rate of air sparging was held constant at 0.3 L air/min using a mass flow meter to ensure a constant CO.sub.2 supply. The reactor is constantly mixed using a magnetic stir bar.

[0122] Experiments were performed at three current densities (6.3, 12.5, and 31.3 A/m.sup.2) corresponding to three different currents (I=0.01 A, 0.02 A, and 0.05 A) to investigate conditions under which the system is kinetically limited by the rate of electrochemical base production (0.01 A), to a condition where base production outpaces alkalinity consumption by gypsum conversion to calcite (0.02 A, 0.05 A). The time evolution of pH and sulfate concentration were monitored throughout the experiment at three sampling ports (SPs) labeled in FIG. 1: the anolyte vessel (SP1), the cathode chamber effluent (SP2), and the effluent of the mixed flow reactor (SP3). Measurements were obtained using a pH electrode. The anolyte was sub-sampled over time for pH and sulfate concentration.

[0123] Solid and Solution Phase Analysis

[0124] Subsampled aliquots of aqueous solution were routinely analyzed for pH measured using a SI Analytics BlueLine pH probe calibrated at pH 1.68, 4, and 10. Solid aliquots were sampled, separated by vacuum filtration, and air dried at ambient temperature for mineralogical characterization.

[0125] Electrochemistry Efficiency Tests Relevant to Other Embodiments

[0126] Efficiency of the electrochemical cell relevant to various embodiments of the process was determined by generating sulfuric acid in batch mode in solutions of sodium, calcium, and magnesium sulfate and mixtures thereof, while intermittently dosing the catholyte with sulfuric acid solution to neutralize the pH (e.g., to remove OH.sup.? by H.sup.++OH.fwdarw.H.sub.2O), thereby mimicking the full process performance. Faradaic efficiency was also calculated for each full-process experiment based on the rate of acid generation in the anolyte solution for a given current.

[0127] Kinetic Modeling

[0128] The dynamic evolution of pH in the mixed flow reactor can be completely described by contributions from three processes: OH.sup.? production in the cathode (f.sub.c, mol OH.sup.?/min), OH.sup.? loss by migration through the AEM (f.sub.AEM), and OH.sup.? consumption by mineral carbonation. Net OH.sup.? produced at the cathode (f.sub.net,c=f.sub.c?f.sub.AEM) enters the mixed flow reactor, and a mass balance on pH in the reactor can be written,

d[OH.sup.?].sub.MFR/dt=f.sub.in?f.sub.out?2R.sub.carb,

[0129] where f.sub.in and f.sub.out represent the flow of OH.sup.? into the mixed flow reactor by introducing catholyte and by removing reactor effluent, respectively, and R.sub.carb is the rate of calcium carbonate precipitation (mol CaCO.sub.3/min). For a volumetric flow rate through the mixed flow reactor v.sub.c (mL/min), we determine the flux of OH.sup.? through the mixed flow reactors at time t based on the pH of catholyte (pH.sub.c; fluid sampled from SP2), f.sub.in(t)=v.sub.c10.sup.?(14?pHc(t)). Similarly, the flux of OH.sup.? out of the MFR depends on the pH of the fluid effluent (pH.sub.eff; fluid sampled from SP2), f.sub.out(t)=v.sub.c10.sup.?(14?pHeff(t)). The rate of mineral carbonation in the mixed flow reactor is determined as a function of time by rearranging the equation above:

R.sub.carb=?0.5[(10.sup.?(14?pHeff(t2))?10.sup.?(14?pHeff(t1)))/(t.sub.2?t.sub.1)?v.sub.c10.sup.?(14?pHc(t))+v.sub.c10.sup.?(14?pHeff(t))].

[0130] Results and Discussion

[0131] Evolution of fluid chemistry and process kinetics. Time resolved measurements of fluid pH in the recirculating anolyte (SP1), catholyte effluent (SP2), and mixed flow reactor effluent (SP3) are given in FIGS. 2A-C for experiments performed at 0.01 A, 0.02 A, and 0.05 A current in the electrochemical cell. The evolution of pH in the catholyte and mixed flow reactor depend on the competing rates of base production at the cathode, which increases pH, counteracted by base consumption by carbonate mineral precipitation. Higher rates of base production were obtained at higher potentiostat currents, such that the highest current experiments constantly evolved towards a higher pH in the mixed flow reactor throughout the course of each experiment (2-4 hour duration).

[0132] The process kinetics were evaluated by applying mass balance expressions to determine the rates of acid and base production as well as the rate of gypsum carbonation (R.sub.carb). The calculated rate of carbonate mineral precipitation R.sub.carb=4.2?0.6 and 4.3?1.2?10.sup.?6 (mol/min) in the 0.02 and 0.05 A experiments, respectively. Invariant rate with solution pH or base production rate suggests that the process is rate-limited by the rate of CO.sub.2 hydrolysis. To confirm this hypothesis, we calculated the mass flux of CO.sub.2 being introduced to the mixed flow reactor by air bubbling. All experiments were performed using a constant volumetric flow rate of 0.29 L/min, which equates to 5.04?10.sup.?6 mol CO.sub.2/min assuming a CO.sub.2 concentration of 380 ppmv. The measured carbonate precipitation rates for the higher current experiments are similar to this CO.sub.2 flux, which supports the conclusion that CO.sub.2 supply is rate-limiting.

[0133] In contrast, for the experiment run at the lowest current (0.01 A), the pH on the cathode side evolved towards a low value that fluctuated from pH?8 to 9.5 (FIGS. 2A-C). These fluctuations can be explained by the observed linear dependence of the carbonation rate on the pH in the mixed flow reactor (FIG. 3). The slope of 0.5 is consistent with the stoichiometry of the carbonation reaction (reaction 2), where 2 moles of OH.sup.? are consumed per mole of carbonate precipitated, indicating that the process kinetics are completely controlled by the rate of base production in this experiment. The process in this case is the most efficient electrochemically, because minimal OH.sup.? is lost across the AEM, but the overall measured rate of carbonation was ?4 times slower in this case compared to the higher current experiments. To optimize the overall process efficiency, both the electrochemical efficiency and the rates of carbonation must be considered. Maximum efficiency will be achieved when the CO.sub.2 hydrolysis flux is equivalent to the rate of base supply, such that the rate of carbonate precipitation is maximized while maintaining a relatively low steady-state pH in the cathode chamber.

[0134] Solid products. The mineralogy of the solid product was determined for an experiment run for an extended duration (10 hours total) at a constant 0.02 A current. The product was demonstrated to contain appreciable calcium carbonate mineral by a fizz test with 10% nitric acid. Fourier Transform Infrared Spectroscopy (FTIR) was used to identify the carbonate product as aragonite, which is a polymorph of CaCO.sub.3.

[0135] Electrochemistry rates & efficiency. The energy consumption required for acid and base generation in the electrochemical cell was measured for the system configuration during individual experiments as well as in batch mode. During the continuous flow process experiments, both the time-averaged rate of acid production and the energy consumption per mole of H.sub.2SO.sub.4 scaled linearly with current (FIG. 4A-B) in solutions of all compositions. Observed values of Faradaic efficiency varied from 48-58% in these experiments and did not vary systematically with the applied current. Higher Faradaic efficiencies were observed in concentrations with higher sulfate concentration in the catholyte (FIG. 4B), due to reduced Voltages as well as lower Faradaic losses from maintaining a higher ratio of sulfate to hydroxide ions in the catholyte solution.

[0136] In batch mode with 1M Na.sub.2SO.sub.4 initial catholyte solution, the Faraday efficiency increases at the beginning of the experiment as the ionic strength in the anolyte goes from zero (DI water is used as the initial solution) to some threshold value (FIG. 5A). The Faraday efficiency of acid generation was high without significant optimization of the electrochemical cell with an average value of 76.8+?0.8% and is independent of the SO.sub.4.sup.2?/OH.sup.? for a wide range of ratios in high sulfate concentrations. To explore the influence of competing OH.sup.? transport across the AEM, we measured the dependence of the Faradaic efficiency on the ratio of aqueous [SO.sub.4.sup.2?]/[OH.sup.?] and found that the efficiency decreases for [SO.sub.4.sup.2?]/[OH.sup.?] ratios less than ?1000 (FIG. 5B). For a system initiated with sulfate and calcium concentrations at stoichiometric saturation with gypsum (0.015 M SO.sub.4.sup.2? based on calculations using PHREEQc geochemical speciation software), this ratio is equivalent to a pH<=9. Thus, if the mixed flow reactor conditions are held at pH<9 by adjusting the rate of base production at the cathode, the process can be maintained at maximum electrochemical efficiency for long durations. Alternatively, additional sulfate can be supplied to the catholyte side of the system (e.g. as sodium sulfate) to maintain a high sulfate concentration relative to hydroxide in the catholyte in more concentrated solutions.

[0137] The energy consumption data combined with the rate of acid production in the batch mode can be used to estimate the energy required per tonne of sulfuric acid produced. We find that the energy intensity varies within the range of 0.13-0.4 kWh/mol H.sub.2SO.sub.4, which is a very low cost compared to typical electrochemical acid/base production and is independent of current density at sufficiently high sulfate concentrations. Energy efficiencies of sulfuric acid and base production by the process described here are on par with the industry-leading chlor-alkali process. The process can achieve similar efficiencies to chlor-alkali by minimizing Faradaic losses. Faradaic losses are avoided in this system by maintaining a high sulfate to hydroxide ratio in the catholyte (>10), which is accomplished in this process by circulating separate solutions through the anode and cathode chambers. At these efficiencies, the unit economics of sulfuric acid production by electrochemical processes becomes economically viable.

[0138] Applications.

[0139] Production of Carbon Negative Sulfuric Acid, e.g., for Critical Element Extraction from Silicate Materials.

[0140] Sulfuric acid leaching and weathering of silicate minerals can neutralize the acidity while at the same time liberating valuable elements. For example, production of lithium carbonates from lithium-bearing claystones (e.g. hectorite, a type of smectite clay mineral) is being explored at large deposits in Nevada. Lithium is commonly extracted from ores using low concentration sulfuric acid leaching (cf. Meshram et al., 2014). Lithium carbonate can be recovered from claystone by reaction with sulfuric acid, for example:

3H.sub.2SO.sub.4+Na.sub.0.3(Mg.sub.2.7Li.sub.0.3)Si.sub.4O.sub.10(OH).sub.2(hectorite)+0.3NaHCO.sub.3.fwdarw.0.15Li.sub.2CO.sub.3+4SiO.sub.2+4.15H.sub.2O+2.7MgSO.sub.4+0.3Na.sub.2SO.sub.4+0.15CO.sub.2

[0141] Completion of the weathering process through protonation and hydrolysis of silicate mineral ores to completely neutralize acidity can be slower than the extraction step. For example, one study showed that smectite minerals can consume and neutralize approximately 1 kg H.sub.2SO.sub.4 per ton of clay mineral per day at 25? C. (Bibi et al., 2014). Some silicate minerals such as serpentine and olivine neutralize acidity much more quickly (McCutcheon et al., 2015; Hamilton et al., 2020).

[0142] In an embodiment (FIG. 6), a process is provided for sulfuric acid leaching of lithium claystone, with return of the magnesium sulfate-containing leachate post-lithium extraction to the mixed flow reactor, recycling the sulfuric acid and precipitating the magnesium sulfate as magnesium carbonate.

[0143] Recycling of Phosphogypsum for Carbon Negative Phosphate Fertilizer Production.

[0144] The main industrial process that consumes sulfuric acid is the production of phosphoric acid for agricultural fertilizer (King & Moats, 2013), which yields gypsum as a byproduct,

Ca.sub.5F(PO.sub.4).sub.3(apatite)+5H.sub.2SO.sub.4+10H.sub.2O.fwdarw.3H.sub.3PO.sub.4+5CaSO.sub.5.Math.2H.sub.2O(gypsum)+HF.

[0145] Other uses of sulfuric acid include mining and extraction of valuable metals such as nickel, copper, and lithium. These PG piles have longstanding environmental and ecological impacts along the global coasts including Florida, Morocco, etc., wherever phosphorus fertilizer is produced. Phosphogypsum deposits also contain elevated concentrations of radionuclides and therefore have limited industrial uses in the United States. However, many of the trace element constituents of PG, including rare earth elements (REEs), uranium, and other metals, have significant value and could be valorized following selective recovery from PG (Mattila et al., 2015; Tayibi et al., 2009).

[0146] In another aspect (FIG. 7), a process is provided for generation of phosphoric acid from rock phosphorus with calcite as the solid product instead of phosphogypsum, for example:

Ca.sub.5F(PO.sub.4).sub.3(fluorapatite)+5CO.sub.2(g)+5H.sub.2O(l).fwdarw.5CaCO.sub.3(calcite)+3H.sub.3PO.sub.4+HF

[0147] The method includes the cyclic steps of electrochemical production of sulfuric acid at the anode and calcium hydroxide aqueous solution at the cathode. The hydroxide solution is reacted with carbon dioxide to produce solid calcium carbonate. In this aspect, the sulfuric acid anolyte is recovered, concentrated as necessary to 93-98% H.sub.2SO.sub.4 and reacted with rock phosphorus to produce phosphoric acid, calcium sulfate, and HF (as in reaction 2). The product calcium sulfate is returned to the process to produce calcium carbonate (as calcite, aragonite, and/or vaterite) and sulfate solution. The sulfate solution is returned to the electrochemical cell along with water to continue the cycle (FIG. 7).

[0148] Further flows charts of steps of embodiments of the invention are shown in FIGS. 8-11.

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