Energy efficient electrodialysis enabled by symmetric reaction couples for lithium extraction and alkalization

Abstract

Systems and methods of lithium extraction and seawater alkalization are provided herein. High operating voltages in electrodialytic processes resulting from large thermodynamic potential and kinetic overpotentials can be greatly reduced by applying a symmetric reaction couple on cathode and anode electrodes. The hydrogen evolution reaction (HER) and hydrogen oxidation reaction (HOR) are demonstrated as the symmetric reaction couple with minimal thermodynamic potential, and ion flow between the cathodic and anodic chambers is controlled by ion-selective membranes. The hydroxide ions are enriched in the cathodic chamber for application of alkalization, where anions such as chloride ions are transported to the anodic chamber for charge balance. In the case of lithium extraction, the lithium ions are transported from the anodic chamber through a lithium-selective membrane and enriched in the cathodic chamber. This technique reduces the cell voltage necessary for driving electrodialytic processes, so that they can become more commercially viable.

Description

PATENT Attorney Docket No.: 079445-1518392-01611 OPC Client Ref No.: S24-329ENERGY EFFICIENT ELECTRODIALYSIS ENABLED BY SYMMETRIC REACTION COUPLES FOR LITHIUM EXTRACTION AND ALKALIZATION CROSS-REFERENCES TO RELATED APPLICATIONS

[0001] This application claims the benefit of priority of U. S. Provisional Application No.63 / 685,151 filed on August 20, 2024, the entire contents of which are incorporated herein by reference.FIELD OF ART

[0002] The present invention pertains to ion-transfer cells, in particular, ion-transfer cells for lithium extraction and seawater alkalization.BACKGROUND

[0003] The rapid growth of lithium-ion batteries has catalyzed an unprecedented demand for lithium. However, global lithium supplies struggle to meet the ever-increasing demand because traditional lithium mining processes are slow, expensive, inefficient, and environmentally unsustainable. The ongoing global warming demands the development of large-scale and sustainable carbon capture techniques, while the current carbon capture technologies like amine scrubbing and calcium looping are heavily depended on energy-intensive and centralized thermal cycling processes.

[0004] In recent years, electrodialytic processes have seen increasing application in various industries, from seawater desalination, wastewater treatment, to lithium extraction and carbon dioxide capture. However, the cell voltage is typically higher than 3 V, which increases the energy consumption of these processes and hinders their practical application.

[0005] Traditional lithium mining processes are slow, expensive, inefficient, and environmentally unsustainable. Thus, remains a need for improved methods and systems for lithium extraction that are more efficient, cost effective and sustainable.BRIEF SUMMARY

[0006] In one aspect, the invention pertains to ion-transfer cells that couple hydrogen evolution reaction (HER) and hydrogen oxidation reaction (HOR) in opposite directions, thereby increasing efficiency of the reactions and allowing for continuous operation at considerably lower voltage, in particular, ion-transfer cells for lithium extraction and seawater alkalization.

[0007] In some embodiments, the high operating voltages in electrodialytic processes resulting from the large thermodynamic potential and kinetic overpotentials can be greatly reduced by applying a symmetric reaction couple on the cathode and anode electrodes. Here, the HER and HOR are demonstrated as the symmetric reaction couple with a minimal thermodynamic potential, and the ion flow between the cathodic and anodic chambers is controlled by ion-selective membranes. The hydroxide ions are enriched in the cathodic chamber for application of alkalization, where anions such as chloride ions are transported to the anodic chamber for charge balance. In the case of lithium extraction, the lithium ions are transported from the anodic chamber through a lithium-selective membrane and enriched in the cathodic chamber. In the case of seawater alkalization, the hydroxide ions are concentrated in the cathodic chamber with chloride ions transported from the cathodic chamber to the anodic chamber through an anion-exchange membrane to balance the charges. This technique reduces the cell voltage necessary for driving electrodialytic processes, so that they can become more commercially viable.

[0008] The demonstrated technique allows ion enrichment through electrodialysis with a significantly lower energy input compared to conventional approaches. In some embodiments, the present invention allows lithium extraction with an energy cost of 1.1 kWh / kg Li, significantly lower than conventional lithium extraction methods. The process is also continuous and exhibits high selectivity, affording sustainable production of high purity lithium hydroxide with low cost. Besides, seawater alkalization could be achieved with 0.063 kWh / kg NaOH, representing only 38% of the theoretical energy consumption (1.64 kWh / kg), and 27% of practical energy consumption (2.34 kWh / kg) through the conventional chlor-alkali process.

[0009] The present invention can be used for energy-efficient electrodialysis applications, including lithium extraction and seawater alkalization processes. The system allowselectricity-driven lithium-mining by separating lithium cations from various other metal cations present in brine water. The source of brine water may come from continental brine, geothermal brine and oil produced water. Also, the system can achieve seawater alkalization with high energy efficiencies, for example, producing sodium hydroxide with lower energy consumption compared to the commercial chlor-alkali process. Similar equipment for selectively extracting other elements may also be developed if suitable materials are available

[0010] In one aspect, the invention pertains to a highly efficient redox-couple electrodialysis (RCE) approach for sustainable lithium extraction from brines. In the current design, the electrodialysis is driven by the same half-cell electrochemical reaction but operated in opposite directions (i.e., the redox couple): hydrogen evolution reaction and hydrogen oxidation reaction, which consume minimum energy due to the theoretically zero-equilibrium full-cell voltage resulting from the nature of the same half-reactions and their low overpotential. Meanwhile, the adoption of solid-state electrolytes as Li-ion selective membranes ensures high selectivity of the lithium extraction from brines with mixed ions. Experiments have demonstrated a continuous lithium extraction from brines for over 100 hours, with a low operating voltage of around 0.25 V, a Faradaic efficiency of 88.87%, and a lithium selectivity of 0.9954. Notably, the lithium extraction via RCE consumes the specific energy of a mere 1.1 kWh per kilogram of Li, approximately an order of magnitude lower than the energy demands of reported lithium extraction techniques. Techno-economic analysis reveals that lithium extraction via RCE offers a substantially reduced cost compared to the traditional lithium extraction technique. The efficiencies and cost benefits innate to the RCE approach described herein not only position it as a promising alternative to prevailing lithium extraction techniques but also as a potential catalyst to reshape the lithium supply chain.

[0011] In another aspect, the invention pertains to Direct Lithium extraction (DLE) techniques based on electrodialysis systems, which have emerged as a promising alternative to conventional approaches, enabling continuous lithium extraction from brine and seawater. Here, a highly efficient redox-couple electrodialysis (RCE) approach has been developed to realize sustainable lithium extraction from brines under an ultralow operating voltage, a high Faradic efficiency and a high lithium selectivity. Techno-economic analysis reveals that lithium extraction via RCE offers a substantially reduced cost compared to the traditional lithium extraction technique. These advantages displayed by this RCE approach overconventional lithium extraction techniques enhance its feasibility in eco-friendly and cost- effective lithium production and could broadly impact the industries of electrified transportation and renewable energy storage.

[0012] In some embodiments, the invention pertains to an efficient redox-couple electrodialysis (RCE) approach to realize sustainable lithium extraction from brines.

[0013] In some embodiments, the invention pertains to zero-equilibrium full-cell voltage resulting from the nature of the same half-reactions and their low overpotential.

[0014] In some embodiments, the invention pertains to continuous lithium extraction from brine under an ultralow operating voltage, a high Faradic efficiency, and a high lithium selectivity.

[0015] In one aspect, the invention pertains to ion-transport cells that utilize hydrogen oxidation reaction (HOR) and a hydrogen evolution reaction (HER) coupled in opposite directions, that allows for greatly reduced operating voltages and high efficiency. In some embodiments, these cells can be run continuously, for at least 100 hours or more.

[0016] In another aspect, the invention pertains to seawater alkalization via an energy¬ efficient electrochemical process for CO₂ capture. Electrochemical processes that increase solution alkalinity therefore enhancing CO?, solubility offer energy-efficient and cost- effective CO?, capture strategies. In some embodiments, the industrial chlor-alkali process produces sodium hydroxide (NaOH) from seawater with a large potential of 2.20 V, dictating by the redox potential difference between the two different redox species at the cathode and anode. Such a large potential can be reduced by applying the same redox species at the two electrodes. Here is used a symmetric reaction couple - hydrogen evolution reaction and hydrogen oxidation reaction - for seawater alkalization with an energy requirement of 0.63 kWh / kg NaOH, significantly lower than theoretical requirements of the chlor-alkali process (1.64 kWh / kg NaOH). This approach demonstrates energy-efficient strategy for seawater alkalization and subsequent CO₂ capture.

[0017] In one aspect, the invention pertains to a lithium extraction cell comprising: a first chamber having a brine input and output, wherein the first chamber is configured for performing hydrogen oxidation reaction (HOR) on the brine in the first chamber; a cathode in the first chamber; a second chamber having a solution inlet and outlet, wherein the secondchamber is configured for performing a hydrogen evolution reaction (HER) on the solution in the second chamber; an anode in the second chamber; a Li-ion selective membrane between the first and second chamber that allows transport of li thium ions from the first chamber into the second chamber; wherein the cell is configured such that, when operated, electrodialysis is driven by the same half-cell electrochemical reaction but operated in opposite directions. In some embodiments, the first and second chamber each include a gas portion that are connected to through a conduit such that H₂ produced by the HER in the second chamber is supplied to the HOR reaction performed in the first chamber. In some embodiments, the Li- ion selective membrane is LAGP. In some embodiments, wherein each of the anode and cathode is a gas diffusion electrodes (GDE). In some embodiments, the first and second chambers are connected by a gas conduit to feed H₂ output from the second chamber by the HER into the first chamber for the HOR. In some embodiments, the extraction cell is configured such that the redox couple of HER and HOR, a theoretical equilibrium potential is zero. In some embodiments, the extraction cell configuration operates at an ultralow voltage of about 500 mV or less, typically 200 mV or less, preferably about 100 mV or less. In some embodiments, the operating voltage is below 0.5 V while the Faradaic efficiency is maintained above 80%. In some embodiments, the cell is operable continuously at 100 hours or more.

[0018] In another aspect, the invention pertains to a method of extracting lithium from brine that includes: feeding lithium-rich brine into a first chamber of a lithium extraction cell; feeding a receiving solution into a second chamber of the lithium extraction cell, wherein the first and second chambers are connected via a lithium -selective membrane; and operating an anode in the first chamber and a cathode in the second chamber to perform hydrogen oxidation reaction (HOR) on the brine in the first chamber and performing a hydrogen evolution reaction (HER) on the solution in the second chamber such that electrodialysis is driven by the same half-cell electrochemical reaction but operated in opposite directions and by the redox pair of HER and HOR, extracting lithium from the brine into the solution to obtain a lithium enriched solution in the second chamber; outputting spent brine from the first chamber; and outputting the lithium enriched solution from the second chamber. In some embodiments, the first and second chamber each includes a gas portion that are connected through a conduit, the method further including:feeding H₂ produced by the HER in the second chamber into the first chamber to feed the HOR. In some embodiments, the Li-ion selective membrane is LAGP. In someembodiments, each of the anode and cathode is a gas diffusion electrodes (GDE). In some embodiments, the first and second chambers are connected by a gas conduit to feed H₂ output from the second chamber by the HER into the first chamber for the HOR. In some embodiments, the extraction cell is configured such that the redox couple of HER and HOR, a theoretical equilibrium potential is zero. In some embodiments, the extraction cell operates at an ultralow voltage of about 500 mV or less, 200 mV or less, typically about 100 mV or less. In some embodiments, the cell operates at a voltage below 0.5 V while maintaining the Faradaic efficiency above 80%. In some embodiments, the extraction cell configuration is operable continuously at 100 hours or more.

[0019] In yet another aspect, the invention pertains to seawater alkalization cell for capturing CO₂, including: a first chamber having an inlet and outlet, wherein the first chamber is configured for performing hydrogen evolution reaction (HER) in the first chamber when the cell is operated; a cathode in the first chamber; a second chamber having an inlet and outlet, wherein the second chamber is configured for performing a hydrogen oxidation reaction (HOR) in the second chamber when the cell is operated; an anode in the second chamber; and an anion-exchange membrane (AEM) between the first and second chamber that allows transport of negatively charged ions from the first chamber into the second chamber; wherein the cell is configured such that, when operated, the coupled HOR / HER reactions produces alkalized seawater that is output from the cell outlets for improved CO₂ capture. In some embodiments, the first and second chamber each includes a gas portion that are connected through a conduit such that H₂ produced by the HER in the first chamber is supplied to the HOR reaction performed in the second chamber. In some embodiments, the AEM allows transport of chloride ions from the first chamber into the second chamber. In some embodiments, each of the anode and cathode is a gas diffusion electrodes (GDE). In some embodiments, the first and second chambers are connected by a gas conduit to feed H₂ output from the second chamber by the HER into the first chamber for the HOR.

[0020] In still another aspect, the invention pertains to a method of alkalinizing seawater to enhance CO₂ capture, the method including: feeding seawater into inlets of first and second chambers of an alkalinization cell, the first and second chambers having a cathode and anode respectively; performing hydrogen evolution reaction (HER) in the first chamber when the cell is operated; hydrogen oxidation reaction (HOR) in the second chamber when the cell is operated; transporting negatively charged ions from the first chamber into the second chamber via an anion-exchange membrane (AEM) between the first and second chambersduring the coupled HOR / HER reactions, thereby producing alkalized seawater that is output from an outlet of the first chamber for improved CO₂ capture. In some embodiments, the method further includes supplying H₂ for the HOR in the second chamber from the first chamber via a gas conduit connecting the first and second chamber. In some embodiments, transporting negatively charged ions is transporting chloride ions via the AEM from the first chamber into the second chamber. In some embodiments, each of the anode and cathode is a gas diffusion electrodes (GDE). In some embodiments, the first and second chambers are connected by a gas conduit to feed H₂ output from the second chamber by the HER into the first chamber for the HOR. In some embodiments, the cell operates at an ultralow voltage of about 200 mV or less, 500 mV or less, typically about 700 mV or less. In some embodiments, the cell operates at a voltage around 0.7 V while maintaining the Faradaic efficiency above 70%. In some embodiments, the extraction cell configuration is operable continuously at 100 hours or more.BRIEF DESCRIPTION OF THE DRAWINGS

[0021] FIG. 1A shows the working mechanism of the continuous lithium extraction via RCE, specifically a schematic of the flow cell designed for continuous lithium extraction from feeding lithium rich brine to the receiving solution, in accordance with some embodiments.

[0022] FIG. 1B shows a comparison between traditional electrochemical DLE methods and this RCE method. Our design employs the reversible redox pair of HER and HOR, offering an equilibrium potential of zero and a substantially reduced polarization potential with increasing current density, in accordance with some embodiments.

[0023] FIG. 2A-2G various aspects of experiments of continuous lithium extraction via RCE, specifically, FIG. 2A shows a photo of the experimental setup, FIG. 2B shows a schematic of the flow cell for electrodialysis, FIG. 2C shows photographs of the LAGP membranes with varied dimensions, in which “D” stands for the diameter, “H” for thickness, and ‘T,” for the side length of a square, FIG. 2D shows SEM images highlighting the cross- sectional microstructure of the LAGP membrane, FIG. 2E shows XRD patterns of the pristine LAGP powder and the sintered LAGP membrane, FIG. 2F shows EIS spectra of the LAGP membrane (upper panel) and the flow cell (lower panel), and FIG. 2G shows a resistanceanalysis for the flow cell with the LAGP membranes of different thicknesses and with varied lithium solution concentrations, in accordance with some embodiments.

[0024] FIGS. 3A-3H shows continuous lithium extraction from Li-rich brine, specifically, time-evolving concentrations of lithium and Na in the brine (FIG. 3 A) and the receiving solution (FIG, 3B) during lithium extraction, FIG. 3C shows Voltage, Faradaic efficiency, and li thium selectivity of continuous lithium extraction, FIG. 3D shows Faradaic efficiency and lithium selectivity of the lithium extraction from the brine with different initial lithium concentrations, FIGS. 3E-3F shows Faradaic efficiency and lithium selectivity of the lithium extraction under different applied current densities (E) and brine flow rates (F) and FIGS. 3G-3H show a summary of operating voltage - Faradaic efficiency plot (FIG. 3G) and energy consumption plot (FIG. 3H) for the RCE approach and the reported electrochemical DLE prototypes, in accordance with some embodiments.

[0025] FIGS. 4A-4B demonstrate long-term stability of the continuous lithium extraction, specifically, the voltage, Faradaic efficiency, and lithium selectivity during the long-term continuous lithium extraction from the brine with mixed Li / Na ions (FIG. 4A) and Li / K ions (FIG. 4B), in accordance with some embodiments.

[0026] FIGS. 5A-5J show various aspects of continuous lithium extraction, specifically, FIG.5A-5D show' SEM images of the LAGP membranes pre (FIG. 5A) and post (FIG. 5B-5D) the 20-hour lithium extraction, FIGS. 5A-5C being is top view of the LAGP surface in contact with the brine, and FIG. 5D is the cross-section view with the upper surface in contact with the brine, FIGS. 5E-5F show XRD patterns (FIG. 5E) and XPS spectra (FIG, 5F) of the LAGP operating for 20-hour lithium extraction under the low' (0.24 mA cm”2) and high current densities (0.96 mA cm”2), where both the XRD patterns and XPS spectra were taken at the surface of the lithium Aluminum Germanium Phosphate (LAGP) membrane m contact with brine, FIG. 5G show's the element ratio deduced from the XPS spectra, FIG. 5H shows the numerical simulation visualizing the lithium distribution in the flow cell during the lithium extraction at varied current densities, FIG. 51 show's the evolution of lithium concentration in the brine side during the lithium extraction under varied current densities, and FIG. 5 J shows a schematic illustrating the lithium depletion near the brine / LAGP interface under a high current density, in accordance with some embodiments.

[0027] FIGS. 6A-6B show techno-economic analyses for the continuous lithium extraction via RCE, specifically, a full process train for the continuous lithium extraction from the brine viaRCE (FIG. 6A) and a techno-economic analysis for two RCE routes, compared against the conventional lithium mining through evaporation and chemical precipitation (FIG. 6B), in accordance with some embodiments.

[0028] FIGS. 7A-7B show the HER / HOR coupled system for seawater alkalization with low energy input, where FIG. 7A shows the voltage ranges of electrochemical reactions for possible seawater alkalization, and FIG. 7B shows the schematic design of a seawater alkalization cell 700, a HER / HOR coupled system for seawater alkalization, including the SHE, standard hydrogen electrode: GDE, gas diffusion electrode; AEM, anion exchange membrane, in accordance with some embodiments.

[0029] FIGS. 8A-8H show' performances of seawater alkalization, specifically, FIG. 8A shows a schematic of the flow device, FIGS. 8B-8D representative current density - time, cell voltage - time, and pH of electrolyte - time curves for the HER / HOR systems for alkalization performance, and FIG. 8E shows the results from COMSOL simulations depicting the gradients of hy droxide ions concentrations [OH⁻] between the anode (HOR) and the cathode (HER) separated by an AEM under different flow rates at ImA / cm2, and FIGS. 8F-8H show heatmaps and illustrate the pH of catholyte, cell voltage and energy cost for NaOH production under different current densities and flow rates for the GDE, gas diffusion electrode and AEM, anion exchange membrane, in accordance with some e bodiments.

[0030] FIGS. 9A-9E show an evaluation of the HER / HOR system for practical applications, specifically, FIGS. 9A-9C show the pH of catholyte, the breakdowns of cell voltages, and the energy cost for producing NaOH under high current density conditions, and FIGS. 9D and 9E are time dependent current density and pH traces of stability test, respectively, and the error bars represent standard deviations (n = 3), in accordance with some embodiments.

[0031] FIGS. 10A-10G show a demonstration of real seawater alkalization, specifically, FIG.10A show's seawater collected from Half Moon Bay area, FIG. 10B a schematic of the procedure for alkalizing seawater, FIGS. C-E show time dependent current density, voltage, and pH traces, respectively, FIG. 10F a schematic of the CO?, adsorption test, and FIG. 10G traces of CO₂ concentration in exhaust gas captured by raw seawater and alkalized seawater, respectively, by the HER, hydrogen evolution reaction and HOR, hydrogen oxidation reaction, in accordance with some embodiments.

[0032] FIGS. 11A-11E show evaluations of energy costs for CO₂ capture by HER / HOR system, specifically, FIG. 11 A shows a schematic of reactions in the HER / HOR system, FIG.11B shows theoretical energy requirements (E_theo) for NaOH production, FIG. 11C shows theoretical energy costs for CO₂ capture with electricity prices ($eie) of 0.01 (orange), 0.03 (pink), and 0.05 (purple) USD / kWh, respectively, FIG. 1 ID shows relation of costs for CO₂ capture to the cell voltage and FE%, the solid lines indicating the certain energy costs of values showing in the plot, and FIG. HE show’s comparisons of theoretical and practical costs for NaOH production via different methods, in accordance with some embodiments.DETAILED DESCRIPTION

[0033] The following descriptions provide additional details as to the experimental setups and studies of the concepts described herein. It is appreciated that the concepts are not limited to these specific implementation and that various other setups could be realized. I. Introduction

[0034] The global transition towards renewable energy and transportation electrification has ushered in an unprecedented demand for lithium (Li). Predictions suggest a remarkable escalation, with the need for lithium rising from approximately half a million metric tons of lithium carbonate equivalent (LCE) in 2021 to an estimated three to four million metric tons by 2030. Historically, the primary’ sources of lithium supplies have been terrestrial resources, predominantly mined through salt-lake brines and high-grade ores via evaporation and chemical precipitation techniques. However, these traditional extraction methodologies present multifaceted challenges. Primarily, the adaptability of these methods is limited due to their sensitivity to the unique geological structures of deposits and their dependence on specific climatic and weather conditions. Such limitations not only complicate the establishment of new mining facilities but also substantially inflate capital investments, driving up the costs of resultant lithium products. The environmental ramifications of these traditional techniques are also concerning. The extensive chemical processes involved are slow, environmentally detrimental, and thus, at odds with the push towards a greener and more sustainable society that electrified technologies promise.

[0035] In the face of these challenges, DLE techniques, especially those rooted in electrochemical processes such as ion-pumping, electrodialysis, and electrolysis, are emerging as viable contenders for lithium mining owing to their environmental compatibility, high efficiency, procedural simplicity, and adaptability to renewable energy. For example, byleveraging the classic intercalation chemistry used in rechargeable lithium batteries, ion¬ pumping provides an effective strategy for separating lithium ions from other competing ions. Many cathode materials for Li-ion batteries, such as LiFePO4and LiMn2O4inherently prefer to incorporate lithium over other ions due to the enhanced structural stability of the intercalation products, which boosts the selectivity of lithium extraction from mixed-ion sources like brine and seawater. However, the necessary post-extraction processes, such as adsorbent regeneration and lithium purification, preclude the continuous operation of ion-pumping systems, introducing substantial efficiency, economic, and environmental costs.

[0036] Electrochemical DLE based on electrodialysis or electrolysis systems offers the potential for continuous lithium extraction from brine and seawater. However, the practical application of this approach seems challenging, partially due to the limited availability of membranes that combine high lithium selectivity and ionic conductivity. Recent advancements in solid-state lithium battery technology have unveiled a broad spectrum of solid-state electrolytes (SSE) that exhibit commendable lithium conductivity and selectivity. Oxide ceramic electrolytes are particularly important due to their high stability in aqueous solutions. Notable examples include NASICON-type Li1.5Al0.5Ge1.5(PO4)3(LAGP) or Li1.5Al0.5Ti1.5(PO4)3(LATP), and perovskite-type Li0.33La0.56TiO3(LLTO). Several current prototypes that employ SSE as the Li-ion selective membrane have demonstrated significant promise for continuous lithium extraction from brine and seawater with high selectivity. However, these prototypes often operate at high voltages due to the substantial redox potential difference between the two electrodes, making the overall process energy-intensive. Additionally, they experience diminished Faradaic efficiency due to side reactions at both electrodes. These inefficiencies raise the operational cost of DLE, thereby inflating the end- product price. Photoelectrochemical and photothermal effects have been integrated into electrodialysis systems to enable green and sustainable lithium extraction, but the expensive materials and delicate structures hinder the scale-up of these devices. Consequently, there is a pressing need for a practical design that facilitates continuous lithium extraction while ensuring low operational voltages and maintaining high Faradaic efficiency — all in the pursuit of sustainable lithium mining.II. Continuous Lithium Extraction from Brine by Efficient Redox-CoupleElectrodialysis

[0037] Presented herein is a highly efficient redox-couple electrodialysis (RCE) approach to realize sustainable lithium extraction from salt-lake and oil-extraction brines. The same half redox reaction is used here as the driving force for electrodialysis but operated in the opposite direction: hydrogen evolution reaction (HER) and the hydrogen oxidation reaction (HOR). The same half reaction offers the zero equilibrium full cell voltage, and HER / HOR has proven to have very low overpotential under realistic current density. Furthermore, the introduction of a Li-ion selective SSE membrane ensures that only lithium ions can transport through the membrane from the brine to the receiving solution, leading to high selectivity and Faradic efficiency for lithium extraction.

[0038] The working mechanism of lithium extraction via RCE is depicted in FIG. 1A. As shown, the flow cell 100 includes a left chamber 10 and a right chamber 20. In the right chamber 20, a LiOH solution flows in, where water undergoes reduction to produce OH’ and Hz at the anode via HER. (Equation (1)); in the left chamber 10, Li-rich brine 1 is supplied, and the H2produced in the right gas chamber 22 flow's to the left gas chamber 12 and then reacts with the OH’ in the brine to form water at the cathode via HOR (Equation (2)).Anode: 2H? O + 2e‘ - -> 112 ’ 2OIT (1) Cathode: 2Hz + 2OI T 2H2O + 2e" (2)

[0039] Driven by the electric field within the electrolyte, lithium ions migrate from the left chamber 10 to the right chamber 2.0, enriching the LiOH solution therein. The spent brine exits the left chamber 10 and the lithium enriched solution 4 exits right chamber 20. The flowcell includes gas diffusion electrodes (GDE) 11, 21, acting as the anode and cathode. The anode being in left chamber 10 to facilitate the HOR, and the cathode being the right chamber 20 to facilitate the HER. The GDE electrodes facilitate exchange of H2 from the solution / bnne, the H2obtained from the HER in the gas chamber 22 of the right chamber 20 fed through connecting conduit 40 into the gas chamber 12 of the left chamber 10 to feed the HOR. The left and right chambers 10, 20 are separated by a lithium-selection membrane 30 such that lithium in the brine passes through the membrane to produce enriched solution in the right chamber 20.

[0040] A distinguishing feature of the RCE design is its reliance on the reversible redox couple of HER and HOR, with a theoretical equilibrium potential of zero. Meanwhile, as the current density rises for HER or HOR the polarization potential remains significantly low'er than in other reactions, as illustrated in FIG. IB. These unique features of RCE allow lithiumextraction to operate at an ultralow voltage (~100 mV), offering its advantages in energy and cost efficiency. By contrast, several recent DLE prototypes via electrodialysis are anchored in oxygen evolution reaction (OER) or chlorine evolution reaction (CER) paired with HER, but suffer from high thermo-dynamic potential difference between OER (or CER) and HER, as well as the large kinetic overpotentials resulting from sluggish reactions, often leading to operational voltages well above 3.0 V.a. Experimental Setups for the Continuous Lithium Extraction via RCE

[0041] The experimental setup for the continuous lithium extraction via RCE is illustrated in FIG. 2A. As shown, the RCE system 200 consists of three compartments, namely the feed compartment 210, the extraction compartment 220, and the collection compartment 230. Within the feed compartment, two syringe pumps 210a, 210b flow the Li-rich brine and the receiving solution into the extraction compartment for lithium extraction. As shown in FIG.2B, the core component in the extraction compartment is a flow' cell 220, which is constructed by stacking acrylic frames with gas diffusion electrodes (GDE) and SSE membrane (LISICON) to form the gas chambers 253, 257 and liquid chambers 254, 256. The brine solution (enriched and spent) is input / output of the first chamber through brine inlet / outlet 251 and the solution (receiving and enriched) is input / output of the adjoining second chamber through solution inlet / outlet 258, During the reaction, H2 is fed from the second chamber from gas outlet 259 through appropriate piping and into the first chamber through gas inlet 252. The spent brine and enriched solution from the flow cell are collected in the collection compartment 230 for further characterizations.

[0042] In the RCE system, the Li-selective membrane is a crucial element, given its profound impact on lithium transport dynamics. In this work, a NASICON-type LAGP was used as the Li-selective membrane due to its high lithium conductivity and selectivity at room temperature. Most importantly, LAGP presents excellent stability in mixed aqueous solutions containing Li, Na, K, and Mg ions. The thermodynamically stable crystal structure of LAGP specific to lithium prevents ion exchange between the embedded lithium in LAGP and other ions, such as Na and K. This stability serves as the foundation for consistent and long-term lithium extraction. The fabrication of the LAGP membrane incorporates traditional sintering techniques, where LAGP powders are first cold-pressed into a green body and then sintered in the furnace without pressure. This methodology offers scalability for the mass production of LAGP membranes of varying sizes and thicknesses (see FIG. 2C). FIG. 2C shows photographsof the LAGP membranes with varied dimensions, in which “D” stands for the diame ter, “H” for thickness, and “L” for the side length of a square. FIG. 2D shows scanning electron microscope (SEM) images highlighting the cross-sectional microstructure of the LAGP membrane. These cross-sectional SEM images of the sintered LAGP membrane reveal its low porosity, which aligns well with its high relative density of 95.0% measured by the Archimedes method. X-ray diffraction (XRD) pattern validates that the sintered LAGP membrane retains the same crystalline composition as the pristine LAGP powders, as indicated in FIG. 2E, which shows XRD patterns of the pristine LAGP powder and the sintered LAGP membrane. FIG. 2F shows EIS spectra of the LAGP membrane (upper panel) and the flow cell (lower panel). Relying on the electrochemical impedance spectroscopy (EIS) spectra measured from the Au / LAGP / Au blocking cell, the ionic conductivity of LAGP as 0.26 mS cm-1at room temperature was calculated, agreeing well with previous reports (0.2 mS cm”1).

[0043] Before initiating the lithium extraction from brine, the electrochemical stability of the flow cell was evaluated. Following a 20-hour settling / forming period, the EIS spectra of the flow cell exhibit negligible variation (see FIG. 2F, lower panel), suggesting that the cell remains undegraded from the solution leakage or membrane corrosion. The calculated resistance of the flow cell is predominantly shaped by two factors: the lithium transport through the LAGP membrane and through the LiOH solution. By tuning the thickness of the LAGP membrane (e.g., 0.13, 0.30, and 0.60 mm) and the lithium concentration in tire solution (e.g., 0,01, 0.10, and 0.50 M), manipulate the resistance associated with lithium transport can be manipulated through these media (FIG. 2G). FIG. 2G show's a resistance analysis for the flow cell with the LAGP membranes of different thicknesses and with varied lithium solution concentrations, in accordance with some embodiments. It also implies that potential reductions in resistance (or operational voltage) for lithium extraction can be achieved by tuning these two factors.b. Continuous Lithium extraction from Li-rich brine via RCE

[0044] First, the efficacy of the continuous lithium extraction from the brine was assessed, with a lithium concentration of 70 ppm (0.01 M) and Na concentration of 2300 ppm (0.1 M). This composition closely mirrors that in the oil-extraction brine (-100 ppm). The flow' rate of the brine and receiving solution entering the extraction compartment is set as 4 cm h’1, and the applied current density for lithium extraction is fixed at 0.08 mA cm"2. This operating condition would, theoretically, enable the extraction of 75% of lithium from the brine, as indicated bythe black dashed line in FIG. 3A, to enrich the receiving solution (0.01 M Li OH). In practice, the lithium concentration in the brine drops from an initial 70 ppm to a stable state of about 25 ppm after the 2 -hour lithium extraction and remains stable thereafter, as shown in FIG, 3 A. Correspondingly, the lithium ions extracted from the brine enrich the receiving solution, as confirmed by the increased lithium concentration in the receiving side, as shown in FIG. 3B. Notably, Na concentrations on both sides remain relatively stable, emphasizing the high lithium selectivity of the RCE approach so that Na ions are not transferred from the brine to the receiving solution. More importantly, throughout the lithium extraction, the operating voltage is impressively low at ~45 mV, and the Faradaic efficiency and lithium selectivity maintain at a high level of 85.5% and 0.9920 (averaged over the 2-10 hour period), respectively (FIG. 3C).

[0045] Per literature reports. Li-rich brines can contain lithium ions with varying concentrations. For example, the lithium concentration ranges from 180 to 250 ppm in the brine at Clayton Valley (Nevada, USA) and is about 690 ppm in the brine at Salar de Olaroz (Jujuy, Argentina). Herein, the performance of lithium extraction from the Li-rich brines is evaluated with different lithium concentrations (70 ppm, 210 ppm, and 630 ppm). The Na concentration in these brines is maintained at 2600 ppm. The Faradaic efficiency and lithium selectivity during the lithium extraction from these varied-concentration brines are consistently high, ranging between 85-90% and 0.99-0.999, respectively, as shown in FIGS. 3D). This indicates that only a small amount of the electricity (10%-- 15%) is lost due to the potential lithium loss carried by the brine flowing and the side reactions at the electrode / electrolyte interface. In general, a higher lithium concentration enhances the extraction efficiency and selectivity, which aligns well with most results reported in the literature. To explore the feasibility of this electrodialysis method in extracting lithium from the brine with higher Na concentration, the performance of continual lithium extraction was tested from brine with lithium concentration of210 ppm and Na concentration of 23000 ppm. The driving force for extracting lithium from the brine to the supporting solution increases from 0.25-0.30 V to 0.7-0.75 V when the brine present had higher salinities. Despite this increase, the voltage required for lithium extraction remains much lower compared to most electrochemical lithium extraction techniques (3-10 V), as shown in FIG. 3G. Similarly, the Faradaic efficiency and lithium selectivity show slight declines but remain within a reasonable range of 70.52 %, and 98.53%, respectively.

[0046] Additionally, the performance of lithium extraction is evaluated under various operating conditions, specifically focusing on the applied current density and the flow rate of feeding brine. Holding the lithium concentration in the brine at 210 ppm and the flow rate offeeding brine at 4 cm h'1, the Faradaic efficiency and lithium selectivity of the lithium extraction were obtained under three different current densities: 0.16 mA cm"z, 0.24 mA cm-2, and 0.32 mA cm"2. These current densities correspond to the extraction of 50%, 75%, and 100% of the lithium content in the brine, respectively. For current densities of 0.16 mA cm"2and 0.24 mA cm-2, the Faradaic efficiency and lithium selectivity remain relatively constant at about 85% and 0.99, respectively (FIG. 3E). For a current density of 0.32 mA cm"2, a slight decrease to 76% in Faradaic efficiency and 0.92 in lithium selectivity was observed. A similar trend is noticed when increasing the flow rate (FIG. 3F). However, for brine with a lithium concentration of 630 ppm, the current density for lithium extraction can reach up to 2.0 mA / cm2, under which the Faradaic efficiency and lithium selectivity are 74% and 0.982, respectively.

[0047] To illustrate the advantages of this RCE approach, its operating voltage and Faradaic efficiency was compared with previously reported DLE techniques (see FIG. 3G). Generally, the operating voltage for most DLE techniques falls in the range of 2.0-6.0 V, with Faradaic efficiency lower than 70%. However, by employing the redox couple HER / HOR, the lithium extraction via RCE manages to reduce the operating voltage to below 0.5 V while maintaining the Faradaic efficiency above 80%, resulting in an impressively low specific energy consumption. For instance, under 0.48 mA crn“2and 8 cm If1, the lithium extraction from the brine with a lithium concentration of 210 ppm operates at a rate comparable with the most state-of-the-art DLE techniques, while it only consumes specific energy of a mere 1.1 kWh per kilogram of lithium (equivalent to 0.105 kWh per kilogram of LCE), approximately an order of magnitude lower than most reported DLE techniques (FIG. 3H). Such reduced energy consumption significantly enhances the feasibility of implementing this RCE approach in practical large-scale applications.

[0048] The long-term performance of the continuous lithium extraction from the brine- containing mixed ions - specifically, blends of lithium and Na ions and blends of lithium and K ions - is examined. Operating under 4 cm h"1and 0.24 mAh cm2, the lithium extraction from the Li-Na mixed brine presents good stability over 100 hours (see FIG. 4A). The voltage shows a plateau around 0.25 V, and the average Faradaic efficiency and lithium selectivity reach as high as 88.87% and 0.9998, respectively. In parallel, this RCE approach also exhibits surprisingly promising efficacy when applied to the brine with lithium and K mixed ions. Maintaining the same flow rate and current density, the Faradaic efficiency and lithium selectivity of lithium extraction remain commendably high, at 83.77% and 0.9954, respectively(FIG. 4B). More importantly, when the device scale is amplified by four times, the performance of continuous lithium extraction still maintains outstanding where the Faradaic efficiency and lithium selectivity are 92.4% and 0.9940, respectively.c. Stability of LAGP Membranes During Continuous Lithium Extraction

[0049] In this section, the electrochemical stability of the LAGP membrane during lithium extraction via postmortem analysis and numerical simulations is evaluated. Two LAGP membranes after 20-hour lithium extraction are collected and rinsed by DI water for evaluation: one under a low current density and flow rate (0.24 mA cm “2and 4 cm h'1) and the other under a high current density and flow rate (0.96 mA cm"2and 16 cm h’1). After operating under a low current density and flow' rate, the surface of the LAGP membrane in contact with brine exhibits very little change compared to its pristine state (see FIGS, 5A and 5B), However, the post¬ operation LAGP membrane exposed to a higher current density and flow' rate showcases an apparent microstructural evolution only at the surface in contact with brine, where the initially dense-packed LAGP grains lose contact with each other, resulting in the increased porosity at the LAGP surface (see FIGS. 5C and 5D), Such morphological evolution indicates the structural degradation of LAGP membranes after long-term exposure to the brine under aggressive operational conditions.

[0050] The evolution in surface morphology and composition is further supported by XRD and XPS analysis (FIG. 5E-5G). Operation under high current density significantly reduces the LAGP crystallinity, as validated by the weakened XRD intensity and the merging of diffraction peaks, such as (211) and (116) orientations, as shown in FIG, 5E. The XPS spectra for the Na signal (FIG. 5F), combined with the calculated element ratio of Na / P (FIG. 5G), imply the infiltration of Na ions into the LAGP surface, likely through the replacement of lithium ions by the larger Na ions. The volume mismatch associated with the Li-Na ion exchange can induce the structural degradation of LAGP membranes (FIG. 5C), such as microcracks, increased porosity, and reduced crystallinity.

[0051] To better understand the structural evolution of LAGP during lithium extraction, numerical simulations were performed to visualize the dynamic ion transport in the flow cell. FIG. 5H illustrates the lithium distribution within the flow’ cell during the lithium extraction at varied current densities. The left domain represents the cross-sectional view? of the lithium brine chamber, and the right domain is for the receiving solution chamber. Under a lower current density and flow' rate (left panel in FIG. 5H), lithium ions in most parts of the brine chambercan transport through the LAGP membrane to the receiving side, as indicated by the uniform light blue color (low lithium concentration) of the spent brine flowing out at the top boundary. However, under a higher current density (right panel in FIG. 5H), only the lithium ions near the LAGP membrane can be transported to the receiving side, leading to local ion depletion in the brine, as highlighted by the dark blue region.[0052 The temporal evolution of the lithium concentration along the outflow boundary (FIG.51) further shows an increase in the lithium concentration gradient in the brine when the current density and flow rate are amplified. In this context, the lithium ions far from the LAGP membrane cannot traverse through the chamber thickness fast enough before being carried out by the flow. Such a phenomenon leads to a high residual concentration of lithium ions in the brine, which negatively affects the measured Faradic efficiency of the lithium extraction under aggressive operating conditions (FIG. 3F). Meanwhile, to maintain charge balance, Na ions are forced to insert into the LAGP at the lithium depletion region (FIG. 5 J), potentially damaging the structural integrity of LAGP membranes (FIGS. 5C and 5E). Given these insights, it becomes unequivocal that optimizing the flow cell design to balance the mass transport of Li, the flow rate of fluids, and the applied current density is paramount for future improvements in the system stability for practical applications (i.e., further increase of the current densities up to ~10 mA / cm2). For example, by depositing a layer of TiO2(6 nm) on the surface of the LAGP membrane via ALD, it was found that Faradaic efficiency and lithium selectivity increase from 65% to 74% and from 0.951 to 0.979, respectively, under the current densities of 0.96 mA / cm2.d. Techno-Economic Analysis for the Continuous Lithium Extraction via RCE

[0053] A preliminary techno-economic analysis is conducted to demonstrate the economic feasibility of continuous lithium extraction via RCE. The cost to deliver the targeted production Li2CO3is selected as the economic metric, where the route with the lower cost is more economically profitable. Two routes of RCE, R-l and R-2, are designed based on the full process train for the continuous lithium extraction via RCE (FIG. 6A). For RCE R-1, the applied current density and flow rate are set as 0.24mA cm-2and 4 cm h'1, respectively. For RCE R-2, these values are 0.48mA cm’2and 8 cm h"1, respectively. The route of the evaporation pond represents the process train for conventional lithium mining via evaporation and chemical precipitation (Route Evaporation Pond), serving as the benchmark for the techno-economic analysis. The total cost (unit: US dollars per metric ton of LCE) for lithium mining via eachproposed tram is the cumulative sum of capital expenses, raw material costs, labor wages, operating and maintenance (O& M) costs, and utilities. A detailed description of the specific assumptions and costing methodologies for each unit process is provided in Table SI and Supplementary Notes.

[0054] Conventional lithium extraction involves the construction of large-scale solar evaporation ponds to concentrate the lithium in the raw brine. The substantial construction cost (64 $ m-2, as reported by Mickley et al.) significantly escalates both capital and O& M costs for lithium extraction. In contrast, the need for evaporation ponds is eliminated in the RCE process, thereby reducing capital costs from 5087 $ ton’1to 2731 $ ton’1for RCE R-l and further down to 1850 $ ton"1for RCE R-2 (FIG. 6B). Additionally, the lithium extraction via RCE bypasses the requisite for sizable quantities of reagents, which alone contributed to a 36% total cash cost for conventional lithium mining in 2019. As a result, raw material costs drop from 2000 $ ton’1to 200 $ ton'1for RCE-R1 (190 $ ton'1for RCE R-2). Moreover, the lithium extraction via RCE consumes a small amount of energy due to its low operating voltages, thereby avoiding the burden of extra electricity cost (25 $ ton-1for RCE R-1 and 60 $ ton-1for RCE R-2). In conclusion, the RCE routes offer a substantially reduced cost profile (4410 $ ton’1for RCE R- 1 and 3543 $ ton’1for RCE R-2) compared to conventional lithium extraction methods (9119 $ ton’1), making it notably competitive against the prevailing market prices of lithium products (37000 $ ton’1). The costs of lithium extraction could be further reduced when adopting even higher current densities and flow rates, positioning the RCE technique as a promising avenue for efficient and cost-effective lithium production from brines.

[0055] In summary, a highly efficient redox-couple electrodialysis (RCE) approach has been developed to realize sustainable lithium extraction from oil extraction and salt-lake brines. Our design leverages the redox couple - hydrogen evolution reaction and the hydrogen oxidation reaction - as the driving force for electrodialysis, which minimizes consumption energy because of their low overpotential and equilibrium potential. Meanwhile, the integration of solid-state electrolytes serving as Li-ion selective membranes ensures high selectivity when extracting lithium from mixed-ion brines. A continuous lithium extraction from brines for over 100 hours, with an operating voltage of around 0.25 V, an average Faradaic efficiency of 88.87%, and an average lithium selectivity of 0.9954 was demonstrated. Notably, the lithium extraction via RCE only consumes a specific energy of 1.1 kWh per kilogram of Li, approximately an order of magnitude lower than the energy demands of most reported lithium extraction techniques. Preliminary techno-economic analysis reveals that the lithium extractionvia RCE offers a substantially reduced cost (4410 $ ton-1) compared to traditional lithium extraction via evaporation and chemical precipitation (9119 $ ton-1), as the RCE approach eliminates the construction of large-scale solar evaporation ponds and the requisite for substantial quantities of reagents. These advantages displayed by the RCE approach over traditional lithium extraction techniques enhance its feasibility in eco-friendly and cost- effective lithium production, with broad implications for the electrified transportation and renewable energy storage industries,e. Calculation of the Faradaic Efficiency, Lithium Selectivity, and Energy Consumption

[0056] The faradaic efficiency was calculated byZFV∆CLi+Faradaic efficiency = -Qwhere Z is the charge number of lithium ion, Fis the Faraday constant, V is the total volume of feeding solution flowing through the cell, & CLi+ is the variation of lithium ion concentration in the receiving solution, and Q is the charge consumption. The nominal lithium selectivity was calculated by∆CLi+Lithium selectivity1= - - - -(∆CLi++ ∆CNa+)where CNa+ is the variation of Na-ion concentration in the receiving solution. The energy¬ consumption is calculated byQUEnergy consumption = ~~~~ — ~~ —» kiA XV? " T" ’ J fwhere t / is the average applied voltage, and Mu+ is the molecular weight of lithium ions.

[0057] Numerical Simulation. The numerical simulations on the ion transport in the solution and through the LAGP membrane are performed using the Tertiary- Current Distribution with the Nernst-Planck-Electroneutrality model integrated in the COMSOL Multiphysics software. The migration, diffusion, and convection of ions (Lr, Na, K!, OH‘) in the solution are governed by the Nernst-Planck equation,dCj+ v •dtzjDF Cj= −Dj∇Ci+ Cjv −RTwhere C, is the ion concentration, Djis the diffusivity of ions, v is the velocity vector of the solution, Zj is the charge of ions, is the Faraday constant,is the electrolyte potential, F isthe molar gas constant, and T is the temperature. The subscript i represents the different ion species (i = Li⁺, Na⁺, K⁺, OH⁻). The distribution of electrolyte potential in the LAGP is solved by the Laplace equation,∇ · im= 0where imis the current density in the LAGP, amis tire ionic conductivity of LAGP, and <pmis the electrolyte potential in LAGP. The interfaces between the solution and LAGP membrane are described by the continual boundaries with Donnan conditions37,RT CLi.φl− φm= −RT / zFln(C / C)Cmwhere C„, is the lithium concentration in LAGP. The geometries and boundary conditions of the 2D simplified model are set to be consistent with the experimental setups.

[0058] Techno-Economic Analysis. Spreadsheet models are built to perform the techno- economic analysis of the proposed routes. The process train of the evaporation pond represents the processes for conventional lithium mining via evaporation and chemical precipitation, serving as the baseline case for the techno-economic analysis. Two routes of RCE (RCE-Rl and RCE-R2) are established based on the full process train for the continuous lithium extraction from brine (FIG. 6A). The current density and flow rate in the RCE R-l are assumed to be 0.24 mA cm"2and 4 cm h"’!, respectively, while the current density and flow rate in the RCE R-2 are assumed to be 0.48 mA cm "2and 8 cm h’1, respectively. The total cost tor lithium mining (unit: US dollars per metric ton of LCE) through all the proposed trains consists of capital expenses, raw material costs, labor wages, operating and maintenance (O& M) costs, and utilities. The capital cost is calculated using a model developed by previous works38, where the local evaporation rate and the volume of the brine for treatment are used to determine the required size of the evaporation pond and the construction fee. The costs of raw materials are obtained from the market and consumer data 2022 by Statista Inc. The labor fee and O& M cost were calculated by data reported in previous works28. For all routes, it is assumed that lithium extraction plants are located near the brine sources, so the cost of pumping the Li-containing brine and the transportation of products and by-products is ignored. Royalty costs for brine operations are also ignored, as these costs can vary significantly depending on the geographic location and local policies.IL Seawater / Alkalization via Energy-efficient Electrochemical Process for CO₂ capture[0059 Electrochemical pH-swing strategies offer a promising avenue for cost-effective and energy-efficient carbon dioxide (CO2) capture, surpassing the traditional thermally activated processes and humidity-sensitive techniques. The concept of elevating seawater’s alkalinity for scalable CO₂ capture without introducing additional chemical as reactant is particularly intriguing due to its minimal environmental impact. However, current commercial plants like chlor-alkali process or water electrolysis demand high thermodynamic voltages of 2.20 V and 1,23 V, respectively, for production of sodium hydroxide (NaOH) from seawater. These high voltages are attributed to the asymmetric electrochemical reactions, where two completely different reactions take place at the anode and cathode. Here, a symmetric electrochemical system was developed for seawater alkalization based on a highly reversible and identical reaction taking place at the anode and cathode. Hydrogen evolution reaction was utilized at the cathode, where the generated hydrogen is looped to the anode for hydrogen oxidation reaction. Theoretical calculations indicate an impressively low energy requirement ranging from 0.06 – 0.48 kWh / kg NaOH for established pH differences of 1.7 – 13.4. Experimentally, alkalization with an energy consumption of 0.63 kWh / kg NaOH was achieved, which is only 38% of the theoretical energy requirements of the chlor-alkali process (1.64 kWh / kg NaOH). Further tests demonstrated the system’s potential of enduring high current densities (~ 20 mA / cm2) and operating stability over an extended period (>110 hours), showing its potential for future applications. Notably, the CO2 adsorption tests performed with alkalized seawater exhibited remarkably improved CO₂ capture dictated by the production of hydroxide compared to the pristine seawater.

[0060] The ongoing global wanning demands the development of large-scale and sustainable carbon capture techniques. Current carbon capture technologies, such as amine scrubbing and calcium looping, heavily depend on energy-intensive and centralized thermal cycling processes. Alternatively, carbon-capture techniques utilizing solid materials like metal organic frameworks (MOF) and covalent organic frameworks (COF) show promising performance while their selectivity and stability against humidity need further developments. In contrast, electrochemical carbon capture gains increased attention nowadays thanks to its relatively low energy consumption, decentralized operation, and high tolerance of humidity. Typical electrochemical carbon capture relies on redox-active species or pH swing to absorb and release carbon dioxide (CO2) molecules. Redox -active species capture and release CO2 molecules under their reduced or oxidized forms, the throughput is determined by the concentrations of these redox-active species and the capture process needs to be isolated from other oxidizingagents such as commonly existing oxygen. On the other hand, pH-swing methods are established based on different solubilities of the dissolved inorganic carbons (DIG) - including dissolved carbon dioxide, carbonic acid, bicarbonate, and carbonate ions - in solutions with different alkalinity levels, where higher alkalinity correlates to higher solubility of the DIG.

[0061] Direct CO₂ capture via alkalized seawater is propitious since the ocean is the second largest carbon pool on earth that stores 38,400 gigatons of CO2 if no extra chemicals are needed as reactant. The industrial chlor-alkali process can split salt solutions into sodium hydroxide (NaOH) and chlorine gas therefore effectively increasing the alkalinity of solutions, however, such an electrochemical reaction requires a cell voltage exceeding 2.20 V (FIG. 1A). Water electrolysis can also increase the solution’s alkalinity by splitting water into hydrogen and oxygen, while a cell voltage of 1.23 V is still required (FIG. 7A). These intense voltage requirements are coming from the asymmetric electrochemical reactions taken place on the cathode and anode, therefore, coupling a redox reaction with its reverse reaction can theoretically further reduce the cell voltage to nearly zero, such as the reaction couple of oxygen reduction reaction (ORR) and oxygen evolution reaction (OER) regardless of its notoriously 4- electron sluggish kinetics.

[0062] Here, a CO₂ capture technique is reported, in which increasing the alkalinity of seawater through the electrochemical hydrogen evolution reaction (HER) is coupled with its reversed reaction, hydrogen oxidation reaction (HOR) (FIG, 7B), As shown m FIG. 7B, the seawater alkalization cell 700, includes a left chamber 710 and a right chamber 720 that receives inputs of seawater 701, the left and right chambers connected by an Anion Exchange Membrane (AEM) 730, and each chamber having a GDE electrode 711, 721 that act as anodes and cathodes that when operated perform HER (at the cathode) in the left chamber 710 and HOR (at the anode) in the right chamber 720. The chambers are further connected by a gas conduit 740 such that Hz is fed into the HOR in the right gas chamber 722 supplied by H2 produced by the HER in the left gas chamber 721. The HER / HOR coupled system outputs alkalized seawater for CO₂ capture 702, In the HER-HOR coupled system, NaOH accumulates in the catholyte through HER, and the generated hydrogen gas is circulated to the anode for HOR. Such a HER-HOR coupled system is endowed by fast reaction kinetics thanks to its simple 2-elect.ron reaction. The thermodynamic potential to drive the system is primarily related to the pH difference between the two chambers and minimal energy is required when the pH difference is small. Experimentally, the alkalization was conducted under different combinations of current densities (0.8 - 20 mA / cm2) and flow rates of feeding solutions (1.2 -48 ml / hr.) for a small device size of 7 × 7 × 2.5 cm. Remarkably low energy input for the NaOH accumulation (0.63 kWh / kg) was achieved experimentally when the flow rate to current density¬ ratio is around 2, representing only 38% of the theoretical energy consumption (1.64 kWh / kg), and 27% of practical energy consumption (2.34 kWh / kg) through the conventional chlor-alkali process. The system’s stability was also assessed by operating the alkalization process for over 110 hours. Furthermore, the alkalized seawater exhibited a significantly enhanced capacity of CCh adsorption compared to the pristine seawater from the CO₂ capture tests.

[0063] To validate the proposed HER-HOR system for seawater alkalization, a flow device was designed and constructed; the device was composed of a feeding part, a reaction part, and a collecting part (FIG. 8A). In the reaction part, gas diffusion electrodes decorated with platinum on carbon black as effective electrocatalysts (Pt / C GDEs) were used as cathode and anode for demonstration, which separated the middle liquid chamber from the two gas chambers at both ends of the flow device; a commercial-available anion exchange membrane (AEM) that allows the transport of chloride ions further divides the liquid space into cathodic and anodic chambers. The feeding part includes two syringe pumps to deliver the feeding solutions to HER and HOR chambers with constant flow rates, respectively, and the resulting solutions with pH changes were collected for subsequent characterizations. A solution of 0.5 M sodium chloride (NaCl) was utilized as artificial seawater for initial investigations unless noted. As shown in FIG. 8A, the alkalization cell 800 includes seawater inlet 801 to the left chamber 802, GDEs 803, 805 as cathode and anode in left and right chambers, an AEM 804 between the left chamber 802 and right chamber 806, an acidified seawater outlet 807 and seawater inlet 808 at the right chamber 806, control circuitry 809 that controls the cathode and anode to perfonn HER and HOR reactions coupled in opposite directions, and an alkalized seawater outlet 810 from the left chamber.

[0064] Firstly, the system’s electrochemical activities of HER and HOR through cyclic voltammetry characterizations were confirmed. Then, a constant current density of 0.82 mA / cm2was applied to drive the electrochemical reactions (FIG. 8B); the cell voltage between the cathode and the anode was monitored and the average cell voltage was calculated to be 0.67 V (FIG. 8C), Concurrently, the receiving solutions from the cathodic and anodic chambers were collected regularly and their pH values were measured (FIG. 8D). Within an hour, the pH values of the catholyte and anolyte shifted from 6.3 to 11.8 and 1.7, respectively, establishing a pH gradient of 10 that remained stable for six hours. The large pH gradient built under such a low voltage supply suggested an energy requirement of 0.63 kWh / kg for NaOH production.>4significantly lower than the energy consumption demanded by the commercial chlor-alkali process (theoretical value of 1.64 kWh / kg, practical value of 2.34 kWh / kg), despite the underdeveloped Faradaic efficiency (FE%) of -65% due to the loss of generated hydroxide through recombination with protons or trapping on the AEM.

[0065] It is considered that the flow rate (nil / hr) to current density (inA / cm2) ratio (r / jappl) would influence the pH gradient established between the cathodic and anodic chambers, which could be a possible strategy to tune the product concentration for future needs. Therefore, a COMSOL. model was built to simulate the gradient of hydroxide ions (OH-) in the flow device under various r / javpi. ratios (1, 2, and 4), assuming an AEM favoring the tran sport of chloride ions. The simulation results revealed a pattern of OH" concentrated near the HER electrode and gradually decayed with increasing distance to the HER side. Particularly, the OH" concentration in the catholytes decreased as the flow rate increased resulting from more rapid fluxes of feeding seawater (FIG. 8E). Meanwhile, nine different combinations of current densities (jappl.= 0.81, 1.02, or 1.22 mA / cm2) and flow rates (r = 1.2, 2.4, or 4.8 ml / hr.) were tested, therefore establishing different r / japplratios (around 1, 1.5, 2, 2,5, 3, 4, 5 and 6). Relatively small current densities were adopted here to minimize the contribution of ohmic drops to the cell voltage as dictated by the system design. Our experimental results revealed variations of hydroxide concentrations in the catholytes under different r / jappl_ ratios (p value = 0.001 from one-way analysis of variance, one-way ANOVA), where the r / jappi, ratios < 4 showed similar pH values around 11.9 (p = 0.22 from two-tail t-test) compared to a pH value of 11.3 when the r / jappi. ratios are larger than 4 (FIG. 8F). The different pH values under different r / jappl_ ratios depict that the OH" concentration is built on the balance between the accumulation through the electrochemical reaction and the dilution through the flush with new feeding solutions, despite other side reactions that might consume the hydroxide ions. The accumulation likely gained the upper hand when the r / japplratio is < 4 while the dilution dominated with a r / jappratio was equal or greater than 4. These experimental observations align with the previous results obtained from COMSOL simulations. Moreover, tire cell voltages were insignificantly changed among different r / jappratios (p value = 0.852 from one-way ANOVA) with an average value of 0.71 ± 0.03 V (FIG. 8G), where -83% of the cell voltage was to meet the thermodynamic requirements for establishing a pH gradient of 10 and other contributions including ohmic drops and overpotentials were similar. Notably, the energy¬ consumptions for producing NaOH were evaluated and a minimum energy consumptions (0.63kWh / kg) were achieved with a rUqjappiratio of < 4 (FIG. 8H). Despite similar cell voltages across different conditions, the FE% were limited when rUqjappi_ ratios were small (< 2) or large (>== 4) possibly due to the side reactions such as recombination of protons and hydroxide over an extended reaction time (small r!i / applratio) or decreased membrane selectivity under higher pressures (larger rliq / japplratio).a. Tests of High Current Densities and Stability

[0066] To evaluate the HER / HOR system’s potential for future applications, its performance was tested under high current densities and an extended operation time, separately. Based on the previous experimental results, a rliq / jappl.ratio of ~2 was maintained. Three current densities 5, 10, and 20 mA / cm2, were tested following similar procedures as previously described. It was found out that the highest pH of the catholyte were 11.93 ± 0.11, 11.88 ± 0.02, and 11.87 ± 0.04 (n = 3) for current densities of 5, 10, and 20 mA / cm2, respectively (FIG.9A). These pH values indicate that similar amounts of OH were retained (p-value = 0.45 from one-way ANOVA) with a fixed rliqjappitratio, which is consistent with previous observations. Specifically, the cell voltages (Ucdi) were examined (FIG. 9B) through three parts: the thermodynamic potential (UHER / HOR) dominated by the pH difference between the two chambers, the ohmic resistance (ηohmic) quantified wi th the help of impedance measurements, and the rest overpotentials (ηoverpotential). The Uceii values were 0.90 ± 0.09, 1.04 ± 0.002, and 1,38 ± 0.01 V (n = 3) for current densities of 5, 10, and 20 mA / cm2, respectively (FIG, 9B), and the increases of Uceii value with greater current densities is mainly coming from the different values of pohmic (0.15 ± 0.003, 0.19 ± 0.07, and 0.23 ± 0.06 V when Jappi, is 5, 10, and 20 mA / cm2, respectively) despite the similar ohmic resistances among different trials (29 ± 1, n = 9). So, tire Uceii under high current densities is heavily affected by the ohmic resistance and can be greatly decreased. Given that the energy cost is proportional to the cell voltage and inversely proportional to FE%, the energy costs for producing 1 kg NaOH were evaluated together with the corresponding Uceii and FE% values. Notably, the energy costs were 1.00 ± 0.14, 1.31 ± 0.07, and 1.80 ± 0.15 kWh / kg NaOH (n = 3) under current densities of 5, 10, and 20 mA / cm2, respectively. This increase in energy requirement primarily came from the higher Uceii values, as the FE% remained insignificantly different across the three current densities (p- value = 0.44 from one-way ANOVA), suggesting the feasibility of reducing the energy input further with future system optimizations. For example, a reduced ohmic resistance of < 0.06 Qas reported in chlor-alkali process can decrease the experimental energy input by 42% under a current density of 20 mA / cm2(1.05 ± 0.09 kWh / kg NaOH, n = 3).

[0067] The stability of the HER-HOR coupled system for alkalizing seawater was also tested by operating the system for an extended experimental duration. A relatively small current density (0.26 mA / cm2) was applied given the limited pumping ability' of the syringe pumps used in this study, and a rliq / jappl.ratio of ~2 was adopted. The system can be steadily operated for over 110 hours (FIG. 9D) with an appreciable average Ucellof only 0.75 V, indicating a maintained catalytic behavior of the Pt / C GDEs. The alkalization performance was examined by measuring the pH of the collecting solutions and a pH difference of ~10 was maintained between the catholytes and the anolytes during the experimental duration (FIG.9E). Additionally, the stability of the AEM was examined by comparing its structural change through Fourier transform infrared spectroscopy, and insignificant changes were noticed before and after a long-term operation. All these results affirm the system’s stability for long-term operation, which is a crucial requirement for future applications.b. Real Seawater Alkalization and CO2 Capture

[0068] To further assess the HER / HOR system for practical CO₂ capture, a seawater alkalization process and CO2 adsorption test was conducted with real seawater collected from Half Moon Bay area in California (FIG. 10A). The concentrations of major ions in the seawater were determined via chromatography and mass spectroscopy characterizations. The collected seawater contained 0.49 M chloride (CT) comparable to the artificial seawater used previously, however, extra metal ions like magnesium (0.05 M Mg2’) exist which will lead to salt precipitation with increased alkalinity through being electrochemical inert under the reaction conditions−2.37 V vs. SHE). To prolong the lifetime of the GDE from salt blockage, a switch -current method was adopted whereas the direction of current flow between the t 'o chambers was alternated every 30 minutes (FIG. 10B).

[0069] Initially, an oxidizing current was applied to chamber 2 for HOR and chamber 1 conducted HER. Then, the direction of current flow' was switched tor every 30 minutes therefore switching the HOR and HER reactions between chamber 2 and chamber 1, respectively. The current density applied to the system was about 10 mA / cm2despite its oxidizing or reducing (FIG. IOC). The cell voltage was monitored (FIG. 10D) and found to share a similar oxidizing / reducing trend as the jappl.- time curve with barely time lag, the averaged absolute cell voltage was calculated to be 0.95 V. The resulting solutions from>7chamber 1 (initial HER, purple in FIG. 10E) and chamber 2 (initial HOR, pink in FIG. 10E) were collected and measured every 15 min. The initial pH of the seawater was 7.92, and the pH of the collecting solutions turned to be alkaline (pH ~ 10) or acidic (pH 2.) depending on the reaction type of HER or HOR, respectively, with negligible time delay. The pH gradient may seem smaller than observed in previous tests with artificial seawater under the same experimental condition, it is due to the greater buffer capacity of the real seawater around pH 10 and the unpreventable precipitation of metal ions like 0,05 M Mg2+(Table SI). Overall, the HER-HOR system is capable of alkalizing real seawater with a reasonable cell voltage.

[0070] Furthermore, the CO2 adsorption ability of the alkalized seawater was evaluated by measuring the remaining CO2 in the exhaust gas while purging 1% CO2 gas through the solutions (FIG. 10F). After switching from argon gas to concentrated CO2 gas, the raw seawater (pH = 7.9) adsorbed CO2based on Henry's law as the CO2gas in the headspace increased to 1 % (green curve in FIG. 4G), In contrast, the CO2concentration in the exhaust running through the alkalized seawater (pH = 9.5) increased in an apparently slower manner compared to the raw seawater, showing a significantly enhanced CO2adsorption capacity (lower curve in FIG.10G) due to the increased alkalinity. The difference in CO2 adsorption ability between the alkalized seawater and raw seawater was 0.2 g CO2 per kg seawater, however, such a value has been greatly underestimated. As noted earlier, the presence of 0.05 M Mg2’ in the real seawater will consume the generated hydroxide as it will be precipitated out when the pH value is greater than 9.23. Considering the CO₂ capture ability of the precipitated salt of magnesium hydroxide (Mg(OH)2), an extra CO2adsorption of ~ 2.2 g / kg seawater should be counted as the Mg(OH)2can also effectively adsorb CO2and form MgCO3. Additionally, the acidified seawater can be neutralized with abundant and low-cost olivine which constitute over 50% of the upper Earth mantle. Mixing the acidified seawater with 0.1 wt.% olivine and leaving in ambient conditions led to an increase of the acidified seawater’s pH value from 2 to 5 within just one week, making it suitable for easy disposal.c. Theoretical Limits of the HER / HOR System

[0071] To explore the potential of the designed HER / HOR system for scalable and energyefficient CO₂ capture, its theoretical limits of energy consumption were assessed. Assuming an ideal system with a unity FE% (FIG. 11 A), the energy input of the HER / HOR system primarily depends on the pH gradient built between the HER and HOR chambers (FIG. 11B). Considering two boundary conditions, alkalizing a non-buffered solution of pH 8,1 - the18average pH of the current ocean - to pH 8.2 requires an energy input of 0.06 kWh / kg NaOH (pH difference of 1.7); while maximizing the hydroxide concentration to ~ 0.5 M which is dictated by the concentration of cations m the seawater needs an energy input of 0.48 kWh / kg NaOH (pH difference of 13.4). Tire global weighted average levelized costs of energy (LCOE) are 0.033 USD / kWh and 0.049 USD / kWh for onshore wind projects and utility-solar scale photovoltaic projects, respectively. The boundary energy costs were calculated to be 0.62 – 4.8, 1.86 – 14.43, and 3.09 – 24.05 USD / ton CO2with assumed electricity prices of 1, 3, and 5 cents / kWh (FIG. 11 C) by using the proposed HER / HOR system for CO₂ capture.

[0072] In practical applications, the energy cost of CO₂ capture is influenced by the electrochemical system’s cell voltage and the FE%. The cell voltage accounts for ohmic drops and overpotentials besides the thermodynamic potential, while the FE% is affected by side reactions like recombination of protons and hydroxide (FIG. 11 A). Therefore, the relation of energy costs to the cell voltage and the FE% with an assumed electricity price of 0.01 USD / kWh were further demonstrated (see FIG. 11D), the HER / HOR system demonstrates promising energy costs compared to all other industrial processes of NaOH production (solid gray bars in FIG. 1 IE). Moreover, the practical energy costs achieved in the lab, 0.63 kWh / kg NaOH, is much lower than all other processes (hatched bars in FIG. 1 IE) despite the limited FE% of approximately 65~70%. It was deemed that further improvements to the HER / HOR system, such as reducing the ohmic resistance via system optimizations and improving the FE% through advanced membranes with higher selectivity, can enhance its alkalization performance and further reduce the energy cost.d. Methods of Testing

[0073] Materials and chemicals'. All materials and chemicals were used as received. Sodium chloride (S7653) was purchased from Sigma-Aldrich. The Pt / C decorated gas diffusion electrode (Pt / C GDE) (11060024) and anion exchange membranes (AEM) (Fumasep FAB-PK- 130) were purchased from the Fuel Cell Store. All deionized (DI) water was obtained from a Millipore water purification system.

[0074] Assembly of flow cell and electrochemical alkalization: Die flow cell components were fabricated from the acrylic sheets using the laser cutter Epilog Fusion M2. Die Pt / C GDEs were used as the anode and cathode, the catalyst loading amount is 0.5 mg / cm260% Platinum on Vulcan. The AEM was soaked in 0.5 M NaCI solution overnight before use. The exposed areas of AEM, cathode, and anode are designed as 1 cm2. Die flow cell components wereassembled with rubber gaskets between each component and tightened by stainless steel bolts and nuts. The feeding and receiving solutions were flowed by two syringe pumps (1000-US, SyringeONE) with a controllable flow rate. The feeding solutions for both cathodic and anodic chambers were 0.5 M NaCl unless noted. The hydrogen gas generated in the cathode chamber was looped to the anode chamber with an additional hydrogen supply of 3 seem flow rate controlled by a mass flow controller (Alicat Scientific). Typically, a constant current (0.8 ~ 20 mA / cm2) was applied and the cell voltage was monitored (0.65 ~ 1.4 V), The experimental durations were 6 hours for low current density tests; the experimen tal durations of high current density tests were 3 hours or when the collecting solutions were beyond 40 ml due to the limitation of syringe pumps. Specifically for seawater alkalization, the seawater was obtained from Half Moon Bay (37°35'56,6" N 122°30'07.7" W) and filtered with a 0.2-pm filter unit (ThermoFisher Scientific 565-0020) to remove possible sand residuals and microorganism; the cathodic current and anodic current were switched for every 30 min for seawater alkalization experiments. The electrochemical experiments, including the chronopotentiometry, the cyclic voltammetry, and the impedance measurements, were conducted by a BioLogic potentiostat. The pH measurements were obtained with a pH meter (SevenCompact S220).

[0075] Tests of CO2capture: Typically, 20 g of testing solution was sealed in a 100 ml bottle with a gas inlet immersed below the solution level and a gas outlet far away from the solution level. The testing solution was firstly purged with Ar for at least 30 min, then 1% CO2 gas (Airgas) was constantly purged into the testing solution with a rate of 30 seem controlled by a mass flow controller (Alicat Scientific). The CO2 concentration in the exhaust gas was monitored by a CO2 sensor (Gaslab CM-20421) and recorded by the Gaslab 2.1 software with an interval of 1 second. Stirring was provided through the experiment.

[0076] Characterizations The scanning electron microscope (SEM) images and the elemental mappings were obtained at Stanford Nano Shared Facilities wdth a Thermo Fisher Scientific Apreo S Lo Vac SEM equipped with a 60 SDD EDS detector. The results of attenuated total reflectance were obtained at. Stanford Nano Shared Facilities with a Nicolet 150 FT / IR Spectrometer. Hie ionic concentrations were obtained at the Environmental Measurements Facility' of Stanford University through Inductively' coupled plasma mass spectrometry (ICP- MS) analysis (Thermo Scientific XSERIES 2 Quadrupole) and ion chromatography (IC) analysis (Dionex ICS 6000 Ion Chromatograph).e. Calculations of efficiencies and costs[0077 The FE% of the NaOH production was calculated by:# of hydroxide ionsFE% - --4--^ - - - x 100%w of input electronsThe theoretical energy input tor NaOH generation was calculated by:...,, ^HER / HORQHER / HORhspec[ / per g | - ------- - - -----'lNaOH‘v‘NaOH where PHER / HORIShie thermodynamic potential for driving the reaction which is theoretically determined by the pH difference between the cathode and the anode, QHER / HOR the charges used for HER and HOR, nNa0His the produced moles of NaOH, and M is the molar mass of NaOH. The practical energy input for NaOH generation was calculated by:F x UEspee ll per g]FE% MNa0Hwhere F is the Faraday constant, U is the ceil voltage including the thermodynamic potential, die ohmic drop, and the overpotentials. The energy costs 'ere calculated by introducing the price of electricity to the energy inputs. The energy inputs and energy costs related to CO₂ capture were calculated based on the assumption that each generated hydroxide captures one CO?, molecule. The one-way analysis of variance (one-way ANOVA) was conducted by using the Scipy library (version 1.3,1) in Python. The two-tailed t-test was conducted by using Microsoft Excel (v. 16.83).0078 ] Simulations.’ The numerical simulations on the ion transport in the solution and through the AEM were performed using the Tertiary Current Distribution with the Nemst-Planck- Electroneutrality model integrated in the COMSOL Multi physics software. The migration, diffusion, and convection of ions (OH, Cl, Na+, H ) in the solution are governed by the Nernst-Planck equation.ZiDjFCyfi.1where C, is the ion concentration, Diis the diffusivity of ions, v is the velocity vector of the solution, z,- is the charge of ions, F is die Faraday constant,. is the electrolyte potential, R is the molar gas constant, and T is the temperature. The subscript i represents the different ionspecies (i = OH, CT, Na+, IT). Tire geometries and boundary conditions of the 2D simplified model were set to be consistent with the experimental setups (FIG. SI). The diffusion coefficients of CT, Oi l. Na+, H+in the 0.5 M NaCl solution were 2x10”, 5.27x]0”9, 1.33xl0“9, and 9.31x10” m2 / s, respectively; the diffusion coefficients of CT, OH, Na+, IF through the AEM were 2xl0”9, 5.27×10−11, 1.33 xlO”11, and 9.3 lx 10”11m2 / s, therefore the AEM was assumed to be chloride-selective.

[0079] In summary, a HER-HOR coupled flow device was developed for energy-efficient and environmentally benign CO₂ capture by alkalizing the seawater without introducing extra chemical components. By manipulating the ratios of flow rate to current density (r / jappl.), the pH gradient between the cathodic and anodic chambers can be altered, supported by the COMSOL simulations and experimental results. Our experimental results demonstrate a rapid elevation of the catholyte’s pH with fast kinetics thanks to the two-electron nature of the reaction. Notably, a remarkable energy cost for NaOH production of 0.63 kWh / kg has been achieved, significantly lower than the commercial chlor-alkali process (1.64 kWh / kg in theory and 2.34 kWh / kg in practical for NaOH production). Furthermore, its feasibility for future applications was validated by operating the HER / HOR systems under high current densities (5 - 20 mA / cm2) and extended operation time (>110 hr), separately. Alkalized seawater can adsorb extra CO2 of 2.4 g / kg seawater, accounting for hydroxide ions both in soluble and precipitated forms, compared to the pristine one. Theoretical evaluations further affirm the HER / HOR’ s promise for energy-efficient CO₂ capture as the calculated energy costs are 0.62 - 4.8 USD / ton CO2 with an assumed electricity price of 0.01 USD / kWh. System designs and material developments - including but not limited to reducing the system’s ohmic resistance, optimizing the mass transport and material designs, or improving the selectivity of membranes for chloride / sodium ions - can further lower the energy costs for CO₂ capture by using the proposed HER / HOR system.[0080 In the foregoing specification, the invention is described with reference to specific embodiments thereof, but those skilled in the art will recognize that the invention is not limited thereto. Various features, embodiments and aspects of the above-described invention can be used individually or jointly. Further, the invention can be utilized in any number of environments and applications beyond those described herein without departing from the broader spirit and scope of the specification. The specification and drawings are, accordingly, to be regarded as illustrative rather than restrictive. It will be recognized that the terms “comprising,” “including,” and “having,” as used herein, are specifically intended to be readas open-ended terms of art. As used herein, the terms “about” or “approximately” can refer to +!- 10% of the recited value. Any references to publications, patents, or patent applications are incorporated herein by reference in their entirety for ali purposes.

Claims

WHAT IS CLAIMED IS:

1. A lithium extraction cell comprising:a first chamber having a brine input and output, wherein the first chamber is configured for performing hydrogen oxidation reaction (HOR) on brine in the first chamber;a cathode in the first chamber;a second chamber having a solution inlet and outlet, wherein the second chamber is configured for performing a hydrogen evolution reaction (HER) on solution in the second chamber;an anode in the second chamber; anda Li-ion selective membrane between the first and second chamber that allows transport of lithium ions from the first chamber into the second chamber;wherein the cell is configured such that, when operated, electrodialysis is driven by the same half-cell electrochemical reaction but operated in opposite directions.

2. lire lithium extraction cell of claim 1, wherein the first and second chamber each include a gas portion that are connected to through a conduit such that H₂ produced by the HER in the second chamber is supplied to the HOR reaction performed in the first chamber.

3. The lithium extraction cell of claim 1, wherein the Li-ion selective membrane is LAGP.

4. The lithium extraction cell of claim 1, wherein each of the anode and cathode is a gas diffusion electrodes (GDE).

5. The lithium extraction cell of claim 1, wherein the first and second chambers are connected by a gas conduit to feed Hi output from the second chamber by the HER into the first chamber for the HOR.

6. The lithium extraction cell of claim 1, wherein the extraction cell is configured such that the redox couple of HER and HOR has a theoretical equilibrium potential of zero.

7. The lithium extraction cell of claim 1, further comprising a power input, wherein the cell operates at an ultralow voltage of about 500 mV or less.

8. The lithium extraction cell of claim 7. wherein the ultralow voltage is about 100 mV or less.

9. The lithium extraction cell of claim 1, wherein operating voltage is below 0.5 V while Faradaic efficiency above 80% is maintained.

10. The lithium extraction cell of claim 1, wherein the cell is operable continuously for at least 100 hours.

11. A method of extracting lithium from brine comprising: feeding lithium-rich brine into a first chamber of a lithium extraction cell; feeding a receiving solution into a second chamber of the lithium extraction cell, wherein the first and second chambers are connected via a lithium-selective membrane;operating a cathode in the first chamber and an anode in the second chamber to perform hydrogen oxidation reaction (HOR) on the brine in the first chamber and performing a hydrogen evolution reaction (HER) on the solution in the second chamber such that electrodialysis is driven by the same half-cell electrochemical reaction but operated in opposite directions and by the redox pair of HER and HOR, extracting lithium from the brine into the solution to obtain a lithium enriched solution in the second chamber;outputting spent brine from the first chamber; andoutputting the lithium enriched solution from the second chamber.

12. The method of claim 11, wherein the first and second chamber each include a gas portion that are connected through a conduit, the method further comprising:feeding H₂ produced by the HER in the second chamber into the first chamber to feed the HOR.

13. lire method of claim 11, wherein the Li-ion selective membrane is LAGP.

14. The method of claim 11, wherein each of the anode and cathode is a gas diffusion electrodes (GDE).

15. The method of claim 11, wherein the first and second chambers are connected by a gas conduit to feed H₂ output from the second chamber by the HER into the first chamber for the HOR.

16. The method of claim 11, wherein the extraction cell is configured such that the redox couple of HER and HOR, a theoretical equilibrium potential is zero.

17. The method of claim 11, further comprising:operating the cell at an ultralow voltage of about 500 mV or less.

18. The method of claim 17, wherein the ultralow voltage is about 100 mV or less.

19. The method of claim 11, further comprising:operating voltage to below 0.5 V while maintaining the Faradaic efficiency above 80%.

20. The method of claim 11, further comprising:operating the cell is operable continuously for at least 100 hours.

21. A seawater alkalization cell for capturing CO2 comprising:a first chamber having an inlet and outlet, wherein the first chamber is configured for performing hydrogen evolution reaction (HER) in the first chamber when the cell is operated;a cathode in the first chamber;a second chamber having an inlet and outlet, wherein the second chamber is configured for performing a hydrogen oxidation reaction (HOR) in the second chamber when the cell is operated;an anode in the second chamber;an anion-exchange membrane (AEM) between the first and second chamber that allows transport of negatively charged ions from the first chamber into the second chamber;wherein the cell is configured such that, when operated, the coupled HOR / TIER reactions produce alkalized seawater that is output from the cell outlets for improved CO?, capture.

22. Tire alkalization cell of claim 21, wherein the first and second chamber each include a gas portion that are connected through a conduit such that H₂ produced by the HER in the first chamber is supplied to the HOR reaction performed in the second chamber.

23. The alkalization cell of claim 21, wherein AEM allows transport of chloride ions from the first chamber into the second chamber.

24. The alkalization cell of claim 21, wherein each of the anode and cathode is a gas diffusion electrodes (GDE).

25. The alkalization cell of claim 21, wherein the first and second chambers are connected by a gas conduit to feed H2 output from the second chamber by the HER into the first chamber for the HOR.

26. A method of alkalinizing seawater to enhance CO₂ capture, the method comprising:feeding seawater into inlets of first and second chambers of an alkalinization cell, tire first and second chambers having a cathode and anode respectively;performing hydrogen evolution reaction (HER) in the first chamber when the cell is operated;performing a hydrogen oxidation reaction (HOR) in the second chamber when the cell is operated; andtransporting negatively charged ions from the first chamber into the second chamber via an anion-exchange membrane (AEM) between the first and second chambers during tire coupled HOR / HER reactions, thereby producing alkalized seawater that is output from an outlet of the first chamber for improved CO₂ capture.

27. The method of claim 26, further comprising:supplying H₂ for the HOR in the second chamber from the first chamber via a gas conduit connecting the first and second chamber.

28. Tire method of claim 26, wherein transporting negatively charged ions comprises transporting chloride ions via the AEM from the first chamber into the second chamber.

29. The method of claim 26, wherein each of the anode and cathode is a gas diffusion electrodes (GDE).

30. The method of claim 26, further comprising:operating the cell continuously for at least 100 hours.

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