ELECTROCHEMICAL CELL AND PROCESS FOR PRODUCING METAL AND A CO-PRODUCT FROM METAL OXIDE AND AN AQUEOUS HALIDE SALT

- University of Oregon

An electrochemical cell and process for producing metal and a co-product from metal ore and an aqueous halide salt are described. The co-product may be a metal hydroxide, halogen, oxygen, and/or a hypohalite. The cell includes a cathode, an anode, and a separator. A catholyte includes (i) water, (ii) a metal hydroxide comprising Q, where Q is an alkali metal, an alkaline earth metal, or a combination thereof, and (iii) suspended metal ore particles comprising MxOy where M is a metal and x and y are integers. An anolyte includes (i) water and (ii) a halide salt comprising Q and X where X is Cl or Br. A process for producing metal includes applying a voltage across the electrochemical cell to effect reduction of the MxOy in the cathode compartment to provide the metal M and a hydroxide comprising Q. X2, O2, and/or XO− is formed in the anode compartment.

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Description
CROSS-REFERENCE TO RELATED APPLICATION

This application is a continuation-in-part of U.S. patent application Ser. No. 18/329,849, filed Jun. 6, 2023, which claims the benefit of the earlier filing date of U.S. Provisional Application No. 63/349,953, filed Jun. 7, 2022, each of which is incorporated by reference in its entirety herein.

FIELD

This disclosure concerns an electrochemical cell and a process for producing metal and a co-product from a metal ore and an aqueous halide salt.

BACKGROUND

Iron and steel production accounts for ˜8% of global emissions contributing to climate change. Most emissions associated with iron and steel production result from stoichiometric quantities of CO2 (˜1.5 kg CO2 per kg Fe) produced during reduction of iron oxide ores in blast furnaces (Fan et al., Joule 2021, 5:829-862; Vogle et al., Joule 2021, 5:2646-2662). Carbon-capture costs can be avoided if the thermochemical reduction of iron ore—which consumes coal and natural gas, producing carbon emissions—is replaced with direct electrochemical reduction, particularly with electricity sourced from renewable resources. The direct reduction of iron metal from ore using zero-emissions electricity is a potentially scalable approach to avoid “difficult to decarbonize” sources of emissions. High temperature electrochemical steel production has been proposed, but capital costs for such processes have historically been incompatible with the production of tonnage metals at <$500/t (the incumbent sale price of steel used for construction). Thus, there is a need for an environmentally-friendly process that is more cost-effective.

SUMMARY

Electrochemical cells and electrowinning processes for producing metal from metal ore and a halide salt are disclosed. In some aspects, the process further produces a metal hydroxide. The cells and processes also co-produce oxygen, a halogen, a hypohalite, or any combination thereof. In certain aspects, the method further produces H2, which can be reacted with a halogen to produce an acid.

An embodiment of the disclosed electrochemical cell includes a cathode, an anode, and a separator. In some embodiments, the cathode comprises low-carbon steel, iron, graphite, vitreous carbon, copper, or titanium, and the anode comprises an oxide coating comprising Ru, Pt, Ir, or any combination thereof, on a conducting substrate such as titanium. The separator may comprise a porous composite or a cation-selective membrane. In certain embodiments, the cell includes a cathode, two anodes, and two separators, wherein a separator is between the cathode and each of the anodes.

In any of the foregoing or following embodiments, the electrochemical cell may further include a gas collecting means for collecting gas generated at the anode and/or a gas collecting means for collecting gas generated at the cathode. The electrochemical cell may also include a voltage source electrically connected to the cathode and the anode. In any of the foregoing or following embodiments, the electrochemical cell may further include a magnet operable to be passed over a surface of the cathode.

In any of the foregoing or following embodiments, the electrochemical cell may further include a catholyte and an anolyte. The catholyte may include (i) water, (ii) a metal hydroxide comprising Q, where Q is an alkali metal, an alkaline earth metal, or a combination thereof, and (iii) suspended metal ore particles comprising MxOy where M is a metal and x and y are integers. The anolyte may include water and a halide salt comprising Q and X where X is Cl, Br, or a combination thereof. The anolyte may have a pH such that oxygen, halogen, or hypohalite ions are selectively generated at the anode.

In one implementation, a cell stack includes a number n of electrochemical cells as disclosed herein, a cathode electrical connector connecting cathodes of each of the electrochemical cells in parallel, and an anode electrical connector connecting anodes of each of the electrochemical cells in parallel. In an independent implementation, a cell stack includes (i) a number n of electrochemical cells as disclosed herein and arranged in series, (ii) a number n-1 of conductive bipolar plates, a conductive bipolar plate positioned between each adjacent pair of electrochemical cells, (iii) a cathode electrical connector connected to a cathode of a first electrochemical cell in the series, and (iv) an anode electrical connector connected to an anode of a last electrochemical cell in the series.

An electrowinning process for producing metal may include (a) providing an electrochemical cell as disclosed herein and a voltage source electrically connected to the cathode and the anode, (b) providing a catholyte comprising (i) water, (ii) a metal hydroxide comprising Q, where Q is an alkali metal, an alkaline earth metal, or a combination thereof, and (iii) suspended metal ore particles comprising MxOy where M is a metal and x and y are integers, (c) providing an anolyte comprising water and a halide salt comprising Q and X, and (d) applying a voltage across the electrochemical cell to effect (i) reduction of the MxOy in the catholyte to provide the metal M and additional metal hydroxide comprising Q, and (ii) production of O2, X2, XO−, or any combination thereof in the anolyte. In some embodiments, providing the electrochemical cell further comprises providing a cell stack as disclosed herein.

In any of the foregoing or following embodiments, M may be Fe, Mn, Ni, Cr, Co, Zn, or any combination thereof. In some implementations, the metal ore particles comprise Fe2O3. In any of the foregoing or following embodiments, Q may be Na, K, Li, Rb, Cs, Mg, Ca, or any combination thereof. In some implementations, Q is Na. In certain embodiments, the anolyte comprises concentrated seawater. In any of the foregoing or following embodiments, the catholyte may include from 50 g/L to 500 g/L of the suspended metal ore particles prior to applying the voltage and/or the anolyte may include from 10 wt % to 50 wt % of the halide salt prior to applying the voltage.

In any of the foregoing or following embodiments, the process may further include continuously or periodically removing oxygen and/or halogen generated in the anolyte and periodically removing at least a portion of the metal M from the cathode. In some implementations, the process further includes continuously or periodically removing hydrogen generated in the catholyte. In some embodiments, the process further includes periodically adding a quantity of the metal ore particles to the catholyte and/or periodically adding a quantity of the halide salt to the anolyte. When the metal ore particles are obtained from a metal ore feedstock further comprising aluminates and/or silicates, the process may further include leaching at least a portion of the aluminates and/or silicates from the metal ore feedstock to provide the metal ore particles.

The foregoing and other objects, features, and advantages of the disclosure will become more apparent from the following detailed description, which proceeds with reference to the accompanying figures.

BRIEF DESCRIPTION OF THE DRAWINGS

The patent or application file contains at least one drawing executed in color. Copies of this patent or patent application publication with color drawing(s) will be provided by the Office upon request and payment of the necessary fee.

FIGS. 1A and 1B are schematic diagrams illustrating an acidic iron process (FIG. 1A) and an alkaline iron process (FIG. 1B).

FIGS. 2A-2C are schematic diagrams of an exemplary electrochemical cell (FIG. 2A), a cathode and anode, each comprising an active material and a current collector (FIG. 2B), and another exemplary electrochemical cell with several additional components (FIG. 2C).

FIGS. 3A-3C are partial schematic diagrams of exemplary electrochemical cells with an external magnet (FIG. 3A), a magnet on a back side of the cathode (FIG. 3B), or a magnet between the cathode and the separator (FIG. 3C).

FIG. 4 is a schematic diagram of an exemplary cell stack comprising a plurality of electrochemical cells.

FIG. 5 is a schematic diagram of another exemplary cell stack comprising a plurality of electrochemical cells.

FIGS. 6A-6D are schematic diagrams of exemplary cell stack units comprising two anodes and one cathode.

FIGS. 7A and 7B are exemplary flow diagrams of an exemplary cathode flow plate (FIG. 7A) and an exemplary anode flow plate (FIG. 7B).

FIG. 8 is a flowchart illustrating one exemplary electrowinning process as disclosed herein.

FIG. 9 is a flowchart illustrating another exemplary embodiment of the electrowinning process as disclosed herein.

FIG. 10 is a schematic diagram illustrating aspects of an electrochemical cell used in one exemplary embodiment of an electrowinning process as disclosed herein.

FIG. 11 is a flowchart illustrating yet another exemplary embodiment of the electrowinning process as disclosed herein.

FIG. 12 is a flowchart illustrating still another exemplary embodiment of the electrowinning process as disclosed herein.

FIG. 13 is an exploded schematic diagram of an exemplary electrochemical cell as disclosed herein.

FIG. 14 is a graph showing half-cell measurements for iron plated at a graphite cathode and Cl2/O2 evolved at a Ti/RuO2 anode.

FIGS. 15A and 15B show measured cell potential over time for a full cell with a constant applied current of 100 mA (FIG. 15A), and a microscopy image of iron sponge formed at the cathode along with a small quantity of red Fe2O3 (FIG. 15B).

FIGS. 16A and 16B are scanning electron microscopy (SEM) images of α-Fe2O3 aggregate particles (FIG. 16A) consisting of nanoscale primary particles (FIG. 16B).

FIGS. 17A and 17B are an optical microscope cross-section image of deposited Fe (gray) on a cathode showing adsorbed Fe2O3 (red) on the growth surface (FIG. 17A) and an SEM image of Fe plated at 100 mA cm−2 (FIG. 17B).

FIGS. 18A and 18B are an SEM image showing an as-deposited RuOx surface as an anodic catalyst (FIG. 18A) and an SEM image of a cross-section of the Ti-supported RuOx catalyst layer (FIG. 18B); the cross-section was prepared by milling with a Ga-focused ion beam.

FIGS. 19A-19D show current-voltage behavior of a batch chlorine-iron (chlor-iron) cell. FIG. 19A is a polarization curve recorded after a 30 min break-in period representing steady-state operation of the cell; FIG. 19B shows powder x-ray diffraction of reagent α-Fe2O3 (upper tracing) and the reduced Fe product (lower tracing) after 30 min of electrolysis at 100 mA cm−2, with simulated diffraction peaks along the x-axis for α-Fe (single line at ˜45 degrees) and α-Fe2O3; FIGS. 19C and 19D show SEM images of the Fe surface after 30 min of electrolysis at 100 mA cm−2 (FIG. 19C) and 300 mA cm−2 (FIG. 19D); both inset scale bars represent 20 μm.

FIG. 20 is an SEM image showing nanoparticles and nanosheets observed on a deposited Fe surface.

FIGS. 21A-21D are top-down images of the surface morphology of Fe films as a function of the applied current density J. FIGS. 21A and 21B are optical plan view images of an Fe surface after 30 min of electrolysis at J=100 mA cm−2 (FIG. 21A) and J=300 mA cm−2 (FIG. 21B); FIGS. 21C and 21D are plan view SEM images of the Fe surface after 30 min of electrolysis at J=100 mA cm−2 (FIG. 21C) and J=300 mA cm−2 (FIG. 21D) (inset images were collected at increased magnification of the same surface).

FIGS. 22A-22D show process parameters and sensitivities of a batch chlor-iron cell with a Nafion® 115 membrane separator. FIG. 22A shows the relationship between ηFe and J for Fe plated; FIG. 22B shows the relationship between ηCl2 and J for RuOx anodes, evaluated in the batch cell with 5.7 mol kg−1 NaCl (pH=2) in both the anolyte and catholyte; FIG. 22C shows the relationship between ηFe and SiO2 content in a representative ore mixture containing 0-8 wt % SiO2 with the remainder Fe2O3 (the total amount of solids fed to the reactor was kept constant at 25 wt %). ηFe was measured in an identical cell to FIG. 22A with the J=100 mA cm−2; and FIG. 22D shows the polarization behavior of the batch cell in the presence of SiO2 at 0 wt %, 1 wt %, and 8 wt % (solids basis).

FIGS. 23A and 23B show the current-voltage performance of a chlor-iron cell separated by a porous diaphragm (Zirfon® PERL UTP 500 separator). FIG. 23A shows the polarization response of the cell after 30 min electrolysis at J=75 mA cm−2; FIG. 23B shows the voltage performance and ηFe of the chlor-iron cell (two separate trials) at a constant J=75 mA cm−2.

FIG. 24 compares the polarization behavior of the chlor-iron flow cell of FIGS. 23A-23B after 30 min of electrolysis at J=100 mA cm−2 and 4 hours of electrolysis at J=100 mA cm−2; the maximum energy efficiency was computed assuming unity Faradaic efficiency towards products.

FIGS. 25A-25E shows performance of a continuous flow chlor-iron cell including an N2030 bilayer membrane separator. FIG. 25A shows the steady-state polarization behavior of the chlor-iron cell; FIG. 25B shows cell voltage vs. time for a constant-current electrolysis at 100 mA cm−2 for 4 h with periodic measurements of ηCl2 and an average ηFe assessed at the end of electrolysis; FIGS. 25C-25E are images of a cleaved cross-section (FIG. 25C), a focused-ion-beam cross-section (FIG. 25D), and a top-down image (FIG. 25E) of the film produced in FIG. 25B).

FIGS. 26A and 26B show Ecell-t data for five consecutive electrolysis runs performed in a cell with a bifacial cathode at J=100 mA cm−2, with Fe collected from the cathode in between trials (FIG. 26A) and a photograph of Fe collected during the consecutive electrolysis runs (FIG. 26B).

FIGS. 27A-27C show SEM image of crushed and sieved 150-300 mesh α-Fe2O3 particles (FIG. 27A), <300 mesh low-grade α-Fe2O3 particles (FIG. 27B), and an EDX spectrum of the low-grade α-Fe2O3 particles (FIG. 27C).

FIGS. 28A and 28B show SEM images of the morphology of Fe after electrolysis at J=100 mA cm−2 of crushed, low-grade ore sieved to 150-300 mesh (FIG. 28A) and <300 mesh (FIG. 28B).

FIGS. 29A-29C are an SEM image of the iron oxide particles produced via thermal oxidation of iron (III) nitrate nonahydrate, Fe(NO3)3·9 H2O, and then crushed and sieved (−300 mesh) (FIG. 29A); powder XRD of calcined Fe(NO3)3·9 H2O after crushing and sieving; all diffraction peaks were consistent with α-Fe2O3 (FIG. 29B); and voltage response of the batch cell fed with Fe2O3 particles, resulting in ηFe=61% after 1 hour of electrolysis at J=100 mA cm−2 (FIG. 29C).

FIGS. 30A and 30B show elemental analysis of iron particles produced in an exemplary chlor-iron cell. FIG. 30A is an SEM image of a particle analyzed via energy dispersive X-ray spectroscopy (EDX); FIG. 30B is a raw EDX spectrum used to calculate elemental composition of the particle.

FIGS. 31A-31C are an SEM image of precipitated Fe2O3 having a surface area of 65 m2g−1 (FIG. 31A), an SEM image of annealed Fe2O3 having a surface area of 4.5 m2g−1 (FIG. 31B), and a graph showing partial current density towards Fe production as a function of applied voltage of the particles of FIGS. 31A and 31B as well as particles having 13 m2g−1 surface area (FIG. 31C).

 DETAILED DESCRIPTION

Embodiments of an electrochemical cell and a method of using the electrochemical cell are disclosed. Certain embodiments of the electrochemical cell and process provide direct reduction of metal ore to metal along with co-production of halogen gas and/or oxygen gas from a concentrated aqueous halide solution (e.g., concentrated seawater). Embodiments of the electrochemical cell and process further co-produce a metal hydroxide. In some embodiments, the electrochemical cell and process may further co-produce hydrogen gas.

Previous processes, acidic and alkaline, have disadvantages compared to the embodiments of the disclosed process. An exemplary acidic iron process is shown in FIG. 1A. An acidified solution containing dissolved iron (e.g., a waste stream from a chemical process) in HCl is placed in a cell 10 including a cathode 12 and an anode 14 The cathode 12 may be, e.g., graphite. The anode 14 may be, e.g., titanium mesh. Iron is deposited at the cathode 12 and Cl2 gas is produced at the anode 14 When starting with iron ore, the process requires dissolving the iron ore with HCl and may further include reducing Fe3+ to Fe2+ to prevent corrosion of iron metal during plating. The dissolution step is costly and the presence of both Fe3+ and Fe2+ leads to parasitic currents at both electrodes that reduce the process efficiency. Parasitic currents involve the partial reduction of a dissolved metal species at the cathode:

The reverse reaction occurs at the anode:

An exemplary alkaline iron process is shown in FIG. 1B. An alkaline solution containing iron ore particles (Fe2O3) and NaOH is placed in a cell 10 including a cathode 12 and an anode 14 Iron is deposited at the cathode 12 and O2 gas is produced at the anode 14 The alkaline iron process poses two major problems, particularly for scaled-up processes: (1) the oxygen-rich environment at the anode 14 can mix with trace hydrogen gas leading to an explosive gas environment; and/or (2) the price per ton of oxygen gas is so low that the process margins are entirely dependent on the sale price of the iron.

I. Definitions

The following explanations of terms and abbreviations are provided to better describe the present disclosure and to guide those of ordinary skill in the art in the practice of the present disclosure. As used herein, “comprising” means “including” and the singular forms “a” or “an” or “the” include plural references unless the context clearly dictates otherwise. The term “or” refers to a single element of stated alternative elements or a combination of two or more elements, unless the context clearly indicates otherwise.

Unless explained otherwise, all technical and scientific terms used herein have the same meaning as commonly understood to one of ordinary skill in the art to which this disclosure belongs. Although methods and materials similar or equivalent to those described herein can be used in the practice or testing of the present disclosure, suitable methods and materials are described below. The materials, methods, and examples are illustrative only and not intended to be limiting. Other features of the disclosure are apparent from the following detailed description and the claims.

The disclosure of numerical ranges should be understood as referring to each discrete point within the range, inclusive of endpoints, unless otherwise noted. Unless otherwise indicated, all numbers expressing quantities of components, percentages, temperatures, times, and so forth, as used in the specification or claims are to be understood as being modified by the term “about.” Accordingly, unless otherwise implicitly or explicitly indicated, or unless the context is properly understood by a person of ordinary skill in the art to have a more definitive construction, the numerical parameters set forth are approximations that may depend on the desired properties sought and/or limits of detection under standard test conditions/methods as known to those of ordinary skill in the art. When directly and explicitly distinguishing embodiments from discussed prior art, the embodiment numbers are not approximates unless the word “about” is recited.

Anode: An electrode through which electric charge flows into a polarized electrical device.

Bipolar plate: A plate that provides an electrical connection between two adjacent electrochemical cells, particularly between an anode compartment of one electrochemical cell and a cathode compartment of an adjacent electrochemical cell. The bipolar plate conducts current between the adjacent cells.

Bulk average pH: An average pH within an anolyte or catholyte, in contrast to a localized pH adjacent an electrode or separator or a localized pH within a pH gradient.

Cathode: An electrode through which electric charge flows out of a polarized electrical device.

Cation-selective: The term “cation-selective” refers to a material that allows cations to pass through the material, but inhibits the movement of anions through the material.

Cell: As used herein, a cell refers to an electrochemical device in which a chemical reaction may be induced by an electric current. A cell stack is an electrochemical device including two or more cells.

Composite: A solid material composed of two or more constituent materials having different physical and/or chemical characteristics that, when combined, produce a material in which each substance retains its identity while contributing desirable properties to the whole. By “retains its identity” is meant that the individual materials remain separate and distinct within the composite structure. A composite is not a solid solution or a simple physical mixture of the constituent materials. In other words, each particle of the composite includes regions or domains of the two or more constituent materials.

Current efficiency: As used herein, the current efficiency refers to the fraction, or percentage, of current that drives a desired chemical reaction, for example the current that produces a metal from a corresponding metal oxide, with the remaining current participating in side reactions.

Electrical connector: A conductive connecting device between a voltage source and an electrode. In some implementations, the electrical connector is a wire or filament.

Electrolyte: A substance containing free ions that behaves as an electrically conductive medium. An electrolyte in contact with the anode may be referred to as an anolyte, and an electrolyte in contact with the cathode may be referred to as a catholyte.

Electrowinning: Recovery of metals from solutions by electrolysis. Also used to refer to electrodeposition of metal from an ore.

Hypohalite: An oxyanion containing a halogen X in a +1 oxidation state—XO−.

Low-carbon steel: Steel containing up to 0.30 wt % carbon.

Separator: A separator is a porous sheet or film placed between the anode and cathode. It prevents physical contact between the anode and cathode and the mixing of gases while facilitating ionic transport.

Sparingly soluble: Requiring at least 1000 g of solvent to dissolve 1 g of solute; a solubility less than 1 g/1000 mL.

Vitreous (glassy) carbon: A non-graphitized and non-graphitizable carbon (cannot be converted into crystalline graphite). The chemical structure is made of sp2-bonded carbon atoms. Vitreous carbon is impermeable, resistant to chemical attack, and has a relatively low density (˜1.5 g/cm3 with isolated closed pores).

II. Electrochemical Cells and Cell Stacks

Embodiments of electrochemical cells and cell stacks for use in an electrowinning process are disclosed. In some embodiments, the process is used to co-produce metal and halogen from a catholyte comprising a metal ore and an anolyte comprising a halide salt. Depending on the operating conditions, the net reaction is:

where M is a metal, X is a halide (e.g., Cl or Br) and x and y are integers.

In certain embodiments, the process is used to co-produce metal and oxygen from a catholyte comprising a metal ore and an anolyte comprising a halide salt. The net reaction is:

where M, X, x, and y are as previously defined.

In certain embodiments, the process is used to co-produce metal and a hypohalite from a catholyte comprising a metal ore and an anolyte comprising a halide salt. A hypohalite may form under conditions where X2 is soluble and reactive in the anolyte. The net reaction is:

where M, X, x, and y are as previously defined.

Electrochemical Cells

FIGS. 2A-2C illustrates exemplary embodiments of an electrochemical cell 100A, 100C (FIGS. 2A, 2C). The electrochemical cell 100A (FIG. 2A) includes the primary components of the cell. The electrochemical cell 100C of FIG. 2C includes additional components, which are present in some aspects of the cell. It is to be understood that each additional component shown in FIG. 2C individually may be present or absent in various aspects of the electrochemical cell.

The electrochemical cell 100A, 100C comprises a cathode 110, an anode 120, and a separator 130 between the cathode 110 and the anode 120. In use, a catholyte 140 is present in the cell 100 and is in fluid contact with the cathode 110 and the separator 130, and an anolyte 150 is present in the cell 100 and is in fluid contact with the anode 120 and the separator 130. In some implementations, the separator 130 is mounted between two spacers 132, 134 (FIG. 2C), which permit fluid contact of the separator 130 with the catholyte 140 and anolyte 150. In some arrangements, the cathode 110 is positioned within a cathode compartment 114, which may be defined by a cathode compartment wall 115 and the separator 130, and/or the anode 120 is positioned within an anode compartment 124, which may be defined by an anode compartment wall 125 and the separator 130. In other arrangements, such as when the cathode is a plate cathode, the cathode compartment 114 may be defined by the cathode 110, the separator 130, and portions of the wall 115 extending between the cathode 110 and the separator 130. Similarly when the anode is a plate anode, the anode compartment 124 may be defined by the anode 120, the separator 130, and portions of the wall 125 extending between the anode 120 and the separator 130. In some aspects, a surface 120a of the anode 120 is a distance D from a surface 130a of the separator 130 (FIG. 2A). The distance D may be greater than zero. For example, the distance D may be at least 5 mm, such as from 5 mm to 20 mm. The distance D may depend, at least in part, on the overall size of the cell 100. In certain embodiments, e.g., as shown in FIG. 2C, the electrochemical cell 100C further comprises a voltage source 160 electrically connected to the cathode 110 by a cathode electrical connector 162 and to the anode 120 by an anode electrical connector 164. Electrical current is supplied to the cathode 110 and anode 120, and an ionic current across the catholyte 140, separator 130, and anolyte 150 completes the circuit.

In any of the foregoing or following embodiments, the cathode 110 may comprise low-carbon steel, iron, copper, graphite, vitreous carbon, titanium, or any combination thereof. In some embodiments, the cathode has an active cathode area (i.e., the surface area of face 110a of the cathode in contact with the catholyte) of from 0.2 m2 to 10 m2, such as from 0.2 m2 to 5 m2, 0.2 m2 to 3 m2, 0.2 m2 to 2.5 m2, 0.2 m2 to 2 m2, 0.25 m2 to 2 m2, 1 m2 to 10 m2, 0.5 m2 to 10 m2, or 0.5 m2 to 5 m2. In some implementations, the cathode 110 comprises a cathode active material 111 and a cathode current collector 112, as shown in FIG. 2B. Suitable cathode current collector materials include, but are not limited to, iron, carbon, copper, and titanium. The cathode active material 111 and current collector 112 may be the same material or different materials.

In any of the foregoing or following embodiments, the anode 120 may comprise a metal oxide-coated substrate or current collector, as shown in FIG. 2B. In some implementations, the substrate or current collector is conductive. In some embodiments, the metal oxide coating 121 comprises oxides of Ru, Pt, Ir, or any combination thereof. In some implementations, the anode substrate or current collector 122 comprises a titanium plate, mesh, or foil. In certain embodiments, a mesh substrate facilitates rapid removal of gas (X2 and/or O2) bubbles from the anode surface. In some implementations, the anode is a mixed metal oxide anode comprising a mixture of Ru, Pt, and Ir oxides on a titanium substrate. One exemplary anode is the DSA® electrode (Industrie De Nora S.p.A., Milan, Italy). The anode of the cell may constructed from one or more sheets in electrical contact and have a total active area similar in size to the cathode active area, e.g., 0.2 m2 to 10 m2.

The electrochemical cell 100C may comprise a catholyte mixing means 117 and/or an anode mixing means 127 (FIG. 2C). Any suitable mixing means may be used. For example, the mixing means may be an impeller rotatably attached to a base of the cathode compartment 114 or anode compartment 124, an impeller on a rotating shaft lowered into the cathode compartment 114 or anode compartment 124, or a fluid (electrolyte, compressed air, or inert gas) pumped or recirculated below cathode compartment 114 or anode compartment 124. The mixing means may be constructed of any material that is inert to the catholyte and anolyte, and does not participate in the electrochemical reactions occurring in the electrochemical cell.

The separator 130 is permeable to alkali metal cations and/or alkaline earth metal cations. In any of the foregoing or following embodiments, the separator may be a porous composite with pore diameters between 2 and 200 nm (a diaphragm) or a cation-selective membrane.

In some embodiments, the separator 130 is a porous composite, or diaphragm, comprising a polymer and metal oxide nanoparticle. In one implementation, the porous composite comprises polysulfone and ZrO2 nanoparticles. The composite may be coated onto a support substrate, such as a polyphenylene sulfide fabric. One commercially available polysulfone/ZrO2 porous composite is a Zirfon® PERL separator (Agfa-Gevaert NV, Mortsel, Belgium). The Zirfon® PERL UTP 500 separator has an average pore size of 150±50 nm.

In some embodiments, the separator 130 is a cation-selective membrane. The cation-selective membrane permits passage of alkali metal cations and alkaline earth metal cations. In one implementation, the cation-selective membrane comprises fluorinated polyethylene chains with side groups comprising fluorinated sulfonic acids. One suitable commercially available separator is a Nafion™ sulfonated tetrafluoroethylene-based fluoropolymer-copolymer separator (manufactured by The Chemours Company, Wilmington, Delaware). In some implementations, the separator 130 is a bilayer membrane. Some suitable bilayer membranes include a sulfonated tetrafluoroethylene-based fluoropolymer-copolymer layer and a second layer with a higher ionic strength or higher rejection rate. Exemplary sulfonated tetrafluoroethylene-based fluoropolymer-copolymer cation-selective membranes include, but are not limited to, a Nafion™ N2050 bilayer membrane, a Nafion™ N2030 bilayer membrane, and a Nafion™ 115 membrane, available from Chemours, Wilmington, DE. In another implementation, the cation-selective membrane is a reinforced perfluorosulfonic acid (PFSA membrane), such as a polytetrafluoroethylene (PTFE) substrate (e.g., fabric)-reinforced PFSA membrane (e.g., GI-N417 membrane, available from the Fuel Cell Store, Bryan, TX). In still another implementation, the cation-selective membrane is a hydrogen separator membrane, such as a mesh polyphenylene sulfide substrate (e.g., fabric) coated with a polymer and zirconium oxide, e.g., a ZIRFON PERL UTP-500 membrane (available from Agfa, Mortsel, Belgium).

In any of the foregoing or following embodiments, the electrochemical cell 100C may further comprise gas collecting means for collecting gas (e.g., halogen and/or oxygen gas) generated at the anode 120. In some embodiments, the gas collecting means 170 (FIG. 2C) covers the anode compartment 124 and does not cover the cathode compartment 114. The gas collecting means 170 may comprise an outlet 172 to facilitate collection of gas. Gas may be collected continuously or periodically through the outlet 172 while the electrochemical cell 100C is operating. In any of the foregoing or following embodiments, the gas collecting means 170 may be a gas envelope comprising a rigid or substantially rigid cover that extends from a wall 125 of the anode compartment 124 to the separator 130, the cover further including an outlet 172. In any of the foregoing or following embodiments, the electrochemical cell 100C may further comprise gas collecting means for collecting gas (e.g., hydrogen gas) generated at the cathode 110. In some embodiments, the gas collecting means 174 (FIG. 2C) covers the cathode compartment 114 and does not cover the anode compartment 124. The gas collecting means 174 may comprise an outlet 176 to facilitate collection of hydrogen. Hydrogen may be collected continuously or periodically through the outlet 176 while the electrochemical cell 100C is operating. In any of the foregoing or following embodiments, the gas collecting means 174 may be a gas envelope comprising a rigid or substantially rigid cover that extends from a wall 115 of the cathode compartment 114 to the separator 130, the cover further including an outlet 176. Desirably, the gas envelope or cover is made of a material that does not react with evolved gas and is impermeable to the gas being collected. Suitable materials for the gas envelope or cover include, but are not limited to fiber-reinforced polyvinyl chloride, nickel, and/or nickel-coated carbon steel.

During operation of the electrochemical cell 100A, 100C, a metal M is deposited onto a surface 110a of the cathode 110. In some embodiments, the metal M is magnetic. In some implementations, a removable cathode 110 may periodically be removed from the electrochemical cell 100 and the deposited metal M removed from the cathode surface 110a by any suitable means (e.g., scraping, peeling, or rinsing).

In any of the foregoing or following embodiments, the electrochemical cell 100C may further include a thermal insulation material 190 proximate an external surface of the cathode compartment 114, proximate an external surface of the anode compartment 124, or proximate external surfaces of the cathode compartment 114 and the anode compartment 124 (FIG. 2C). The thermal insulation material 190 may retain heat generated by the electrochemical reactions taking place in the cell 100C and/or may reduce an amount of external energy required to operate the electrochemical cell 100C at a desired temperature. In any of the foregoing or following embodiments, an external heating device (not shown) may be used to raise the temperature within the electrochemical cell to a desired operating temperature prior to using the electrochemical cell. In some embodiments, the external heating device also is used during cell operation to maintain the desired temperature when the electrochemical reactions do not generate sufficient heat to maintain the desired temperature.

In another embodiment (not shown), a single electrochemical cell may include a plurality of spatially separated cathodes in the cathode compartment and a corresponding plurality of spatially separated anodes in the anode compartment. The electrochemical cell further comprises a voltage source electrically connected to the cathodes in parallel by a cathode electrical connector and to the anodes in parallel by an anode electrical connector. Alternatively, a voltage source may be separately connected to each cathode and a corresponding anode by an individual cathode electrical connector and anode electrical connector, respectively.

In any of the foregoing or following aspects, when the catholyte compartment 114 has an open top (e.g., as illustrated in FIG. 2A), which facilitates collection of metal M, chemically inert polymer sheets (e.g., poly(tetrafluoroethylene), PTFE) may be used to prevent leakage of the catholyte. Catholyte leakage may occur due to gassing caused by parasitic currents toward the hydrogen evolution reaction in the catholyte. In some aspects, the polymer sheets (not shown) extend vertically above the catholyte compartment 114 to allow gas bubbles to exit the compartment 114 without causing the catholyte 140 to spill or leak out of the compartment 114.

In some aspects, the electrochemical cell 100C may further comprise a magnet 180 (FIG. 2C). In some implementations, the magnet 180 is a magnetic drawbar or a magnetic rod. The magnet 180 is movable along a length of the cathode 110 to facilitate periodic collection of magnetic metal M deposited onto the cathode surface 110a. In some implementations, as shown in FIG. 2C, the magnet 180 is passed over an opposing surface 110b of the cathode 110. As the magnet 180 is moved along the surface 110b, at least some of the metal M moves along the surface 110a following movement of the magnet 180. As the magnet 180 is lifted out of the electrochemical cell 100, at least a portion of the metal M is transferred from the surface 110a to the magnet 180 and removed from the electrochemical cell 100 with the magnet 180. The magnet 180 may be moved by any suitable means, including manual movement by an operator or automatic movement controlled by an external controller device. In one embodiment, when the cathode 110 is constructed from a magnetic material (e.g., the cathode is iron-based), the electrochemical cell 100 does not include a magnet 180. In an independent embodiment, when the metal ore in the catholyte 140 is magnetic, the electrochemical cell 100A, 100C does not include a magnet 180.

In some implementations, the magnet 180 is immersed in the catholyte 140. In other implementations, the magnet 180 is placed such that the magnet 180 is not immersed in the catholyte 140.

In one embodiment, the cathode 110 is a plate (for example, a graphite plate) such that cathode surface 110b defines an external wall surface of the cathode compartment 114 and the magnet 180 is placed adjacent surface 110b (e.g., as shown in FIG. 2C) such that the magnet 180 is between the cathode surface 110b and the wall 115. In an independent embodiment (FIG. 3A), the cathode 110 is disposed proximate an internal surface of a cathode compartment wall 115 with the magnet 180 proximate an external surface of the cathode compartment wall 115, the wall 115 permitting magnetic interaction between the magnet 180 and deposited metal M. In another independent embodiment (FIG. 3B), the magnet 180 may be a magnetic rod placed proximate a surface 110b of the cathode 110. In such arrangements, the cathode surface 110b may comprise a recessed groove (not shown) to receive the magnetic rod, thereby facilitating minimal spacing between the cathode surface 110b and the wall 115. In the arrangement of FIG. 3B, the magnet 180 may nor may not be immersed in the catholyte 140; for instance, if the cathode 110 is a plate cathode that extends across a width of the cell 100, the magnet 180 may not be immersed in the catholyte 140. In still another independent embodiment (FIG. 3C), the magnet 180 may be a magnetic drawbar or a magnetic rod placed between the cathode 110 and the separator 130.

In any of the foregoing embodiments, the magnet 180 may be actuated and manipulated externally to collect metal deposited on the cathode 110. In some implementations, a portion (dashed line) of the magnetic rod 180 may extend through a lower wall of the cathode compartment 114 to facilitate manipulation of the magnetic rod (FIGS. 3B, 3C).

Electrochemical Cell Stacks

In some embodiments, a plurality of electrochemical cells is assembled to form a cell stack. FIG. 4 is a schematic block diagram illustrating one exemplary embodiment of a cell stack 400. The cell stack 400 includes a number n of electrochemical cells 100 as disclosed herein. The number n is 2 or more. In some embodiments, the number n is from 2 to 100, such as from 5-100, 10-100, 25-100, or 25-75. The cells 100 are arranged in parallel. The cell stack 300 further includes a cathode electrical connector 162 connecting each of the cathodes 110 in parallel, and an anode electrical connector 164 connecting each of the anodes 120 in parallel. The cell stack 400 may further include a voltage source 160 connected to the cathode electrical connector 162 and the anode electrical connector 164. The cells 100 of the stack 400 may be spatially separated as shown in FIG. 4. Alternatively, an electrically nonconductive plate or membrane 410 may be placed between each pair of adjacent cells 100. Exemplary materials for the nonconductive plate or membrane include, but are not limited to, metals (e.g., stainless steel) with an insulating elastomeric coating, non-conducting composites, and non-conducting polymers suitable for use at the stack operating temperature.

FIG. 5 is a schematic block diagram illustrating another exemplary embodiment of a cell stack 500. The cell stack 500 includes a number n of electrochemical cells 100 as disclosed herein. The number n is 2 or more. In some embodiments, the number n is from 2 to 100, such as from 5-75, 5-50, 10-50, or 20-50. The cells 100 are arranged in series. The stack 500 further comprises a number n-1 of conductive bipolar plates 510. A conductive bipolar plate 510 is positioned between each adjacent pair of electrochemical cells 100. Suitable materials for the conductive bipolar plate include, but are not limited to, nickel, nickel-plated steel, conductive carbon (e.g. vitreous carbon or graphite), and conductive composites. In one embodiment, all of the conductive bipolar plates have the same composition. In another embodiment, the conductive bipolar plates may be constructed of different materials. In some implementations, the conductive bipolar plate 510 comprises graphite. In such implementations, the cathode 110 and anode 120 may comprise metal foils on either side of the thermally and electrically conductive graphite, with the anode further comprising an oxide coating as described herein. The cell stack 500 may further include a voltage source 160, a cathode electrical connector 162, and an anode electrical connector 164. The cathode electrical connector 162 is connected to a cathode 110 of the first electrochemical cell 100 in the series. The anode electrical connector 164 is connected to an anode 120 of the last (nth) electrochemical cell 100 in the series. The cell stack 500 may further include a voltage source 160 connected to the cathode electrical connector 162 and the anode electrical connector 164.

The cathode electrical connector 162 and the anode electrical connector 164 may be high-voltage connectors.

In the cell stack 400, 500, each cathode may have the same composition, or at least one cell may include a cathode 110 of a different composition. In the cell stack 400, 500, each anode 120 may have the same composition, or at least one cell may include an anode of a different composition. In the cell stack 400, 500, each separator 130 may have the same composition, or at least one cell may include a separator of a different composition. When prepared for use, the cell stack 400, 500 includes a catholyte 140 in each electrochemical cell 100 (e.g., in the cathode compartment 114) and an anolyte 150 in each electrochemical cell 100 (e.g., in the anode compartment 124). In the cell stack 400, 500, each catholyte may have the same composition, or at least one cell may include a catholyte of a different composition. In the cell stack 400, 500, each anolyte may have the same composition, or at least one cell may include an anolyte of a different composition.

In the cell stack 400, 500, each electrochemical cell 100 may include a gas collecting means (not shown), as previously described, to collect gas generated at the anode. Alternatively, a single gas collecting means may be configured to collect chlorine gas from each of the plurality of electrochemical cells 100. In such an arrangement, the single gas collecting means may be subdivided so that each anode compartment 124 is covered by the gas collecting means while the cathode compartments 114 are not covered by the gas collecting means.

In some embodiments, the cell stack 400, 500 further comprises one or more magnets (not shown) to facilitate collecting deposited metal from the cathodes 110. In one implementation, a magnet is inserted in each electrochemical cell 100. The magnet 180 may be, for example, a magnetic rod (e.g., a polyethylene coated magnet) inserted adjacent a surface of the cathode 110 (e.g., adjacent surface 110b as shown in FIG. 3B). In one arrangement (not shown), the cathode surface 110b may include a recessed groove to receive the inserted magnetic rod. Alternatively, the magnet 180 may be inserted between the cathode 110 and the separator 130 (FIG. 3C). The magnetic rod may be inserted from the top of the cell 100 or through a bottom wall of the cell 100 (e.g., as shown by the dashed portion in FIGS. 3B, 3C). The magnet is actuated and manipulated externally to collect metal deposited on the cathode 110. In another implementation, one or more small permanent magnets may be placed in each cathode compartment 114 and externally manipulated to collect deposited metal. In still another implementation, the cathodes 110 are removable and deposited metal may be removed from the cathode surface by any suitable means (e.g., scraping, peeling, or rinsing).

FIGS. 6A-6C are schematic block diagrams illustrating other exemplary embodiments of an electrochemical cell 600A, 600B, 600C, respectively, suitable for use in a cell stack. In FIGS. 6A and 6B, an electrochemical cell 600A, 600B includes two anodes 120, a cathode 110, two separators 130, an anolyte 150, and a catholyte 140. The cathode 110 is placed between the two anodes 120 with a separator 130 between each anode 120 and the cathode 110. In FIG. 6A, the anodes 120 are spaced apart from the separators 130. In FIG. 6B, the anodes 120 are in direct contact with the separators 130. In FIG. 6C, the cell stack 600C further includes an anode flow plate 620 in each anode compartment 124 and two cathode flow plates 610 in the cathode compartment 114, with one cathode flow plate 610 on each side of the cathode 110. In FIG. 6D, the cell stack 600D further includes an anode flow plate 620 in each anode compartment 124 and two cathode flow plates 610 in the cathode compartment 114, with one cathode flow plate 610 on each side of the cathode 110. In contrast to FIG. 6C, the anodes 120 are in direct contact with the separators 130, and the anode flow plates 620 are in contact with the anodes 120, and may compress the anodes 120 against the separators 130. In FIGS. 6C and 6D, each anode 120 may have the same composition or the two anodes 120 may have different compositions. Similarly, each anolyte 150 may have the same composition or the two anolytes 150 may have different compositions.

In any of the foregoing or following aspects, the cell stack may further include a supportive frame to distribute pressure across a flexible separator 130 between the anolyte and catholyte compartments. In some aspects, the supportive frame is made of a stiff, chemically inert polymer (e.g., high density polyethylene, polyether ether ketone, or polytetrafluoroethylene-coated metal). The open area of the frame may be maximized to maintain ionic conductivity between the compartments and the edges of the frame may be beveled to facilitate gas release and prevent accumulation within pockets of the frame. Exemplary supportive frames are shown in FIG. 6C (see, e.g., cathode flow plates 610), FIGS. 7A-7B (discussed below), and FIG. 13 (elements 1350, discussed below in Example 1).

FIG. 7A shows details of an exemplary cathode flow plate 610. The flow plate 610 includes a support frame 612 and a plurality of apertures or current carrying paths 614 within the support frame 612. The apertures 614 extend through the support frame 612, and allow current to pass through from the cathode to the separator. The flow plate 610 further includes a plurality of integrated flow channels with narrow openings 616, which are used to recirculate the suspension of metal oxide particles within the catholyte or streams of inert gas to purge dissolved O2/H2 from the catholyte. Hydrogen may be generated in the catholyte and oxygen may undesirably diffuse through the separator to reach the catholyte. If catholyte is not recirculated, forced convection of an inert gas (e.g., N2, Ar) through openings 616 may be used to agitate the catholyte and prevent particle settling; such settling reduces selectivity of the current toward metal reduction. The inert gas also may purge dissolved O2 from the catholyte. Dissolved O2 also reduces selectivity of the current toward metal reduction.

FIG. 7B shows details of an exemplary anode flow plate 620. The flow plate 620 includes a support frame 622 and a plurality of channels 624 that allow gas flow through the support frame 622. The flow plate further includes an inlet 626 and an outlet 628 built into the frame for recirculating the anolyte. Typically anolyte is recirculated to the base of the flow plate 620 and excess anolyte may be collected from the top of the anode compartment 124 (FIGS. 6C, 6D). In some aspects, the anode flow plate 620 provides support for an anode 120, such as an anode mesh or expanded metal anode mesh. The flow plate 620 may also be used to compress the anode 120 against the separator 130 in aspects where there is direct contact between the anode 120 and separator 130. The flow plate 620 may include apertures (not shown) that receive fasteners used during cell assembly.

When prepared for use, the electrochemical cell 100, 600A, 600B, 600C, 600D includes a catholyte 140 and an anolyte 150. The catholyte comprises (i) water, (ii) a metal hydroxide comprising Q, where Q is an alkali metal, an alkaline earth metal, or a combination thereof, and (iii) suspended metal ore particles comprising MxOy where M is a metal and x and y are integers. The anolyte comprises water and a halide salt comprising Q and X. Further details of the catholyte and anolyte are discussed below.

III. Process for Producing Metal and Co-Products

The disclosed electrochemical cells and cell stacks may be used in an electrowinning process for producing metal. The process co-produces a metal hydroxide. In some embodiments, the process also co-produces halogen (e.g., Cl2, Br2), oxygen, hypohalite (e.g., ClO−, BrO−), hydrogen, or any combination thereof. In certain embodiments, when a halogen and hydrogen are co-produced, the halogen and hydrogen may be combined external to the electrochemical cell or cell stack to produce an acid HX, e.g., HCl or HBr. When a hypohalite is produced, the hypohalite may react with partially reduced metal oxide to regenerate X− ions. Advantageously, the metal M and co-products are commodity products that are useful in manufacturing processes. For example, the metal hydroxide may be used to leach silicates and/or aluminates from metal ores, as well as in other processes, such as carbon capture.

As shown in FIG. 8, an exemplary electrowinning process 800 includes providing an electrochemical cell as disclosed herein, the electrochemical cell further including a voltage source electrically connected to the cathode and the anode (step 801); providing a catholyte within the cathode compartment, the catholyte comprising (i) water, (ii) a metal hydroxide comprising Q, where Q is an alkali metal, an alkaline earth metal, or a combination thereof, and (iii) suspended metal ore particles comprising MxOy where M is a metal and x and y are integers (step 802); providing an anolyte within the anode compartment, the anolyte comprising water and a halide salt comprising Q and X (step 803); and applying a voltage across the electrochemical cell to effect reduction of the metal ore (MxOy) particles in the cathode compartment to provide the metal M and oxidation of halide ions in the anode compartment to form X2 gas, decomposition of water in the anode compartment to form O2, or production of XO− in the anode compartment (step 804). A metal hydroxide comprising Q is coproduced in the catholyte. In certain aspects, H2 also may be concomitantly formed in the cathode compartment. In some embodiments, the reduced metal M is deposited as a dense film onto a surface of the cathode. The deposited dense film may be a free-standing film. In certain implementations, M is periodically collected as a film, a porous deposit, or fine powder.

The overall reactions at the cathode and anode are as follows:

The net reaction is:

In the above reactions, Q cations provide the balancing charges. Because the separator is permeable to Q cations, as hydroxide anions and X2, O2, and/or XO− are produced, Q cations pass from the anolyte through the separator and into the catholyte to maintain the charge balance. Because the generation of protons and hydroxide occurs in separate catholyte and anolyte streams, this reaction can lead to the net generation of acid and base. In some embodiments, when the predominant product is X2, a small amount of O2 also is generated at the anode, which contributes to an acidic environment at the anode surface and suppresses further generation of O2. The acidic environment facilitates maintaining the X2 in gaseous form in the anode compartment. Cl2 is a gas at temperatures >25° C. An operating temperature >60° C. will provide Br2 in gaseous form. When XO− is a product of the cell, it can either be collected as a chemical byproduct or reacted with a partially reduced metal oxide species (e.g. Fe3O4) to recover X−.

In some aspects, a separation or gap between the anode 120 and separator 130 (see, e.g., FIG. 2A) facilitates a pH gradient within the cell, ensuring that a concentration of protons at the separator surface 130a is lower than a concentration of protons at the anode surface 120a. A low concentration of protons at the separator surface 130a favors Q ions as the primary species carrying current through the separator 130 to the cathode compartment 114. A gap between the anode 120 and separator 130 may be present when oxygen is a desired co-product. As previously set forth, the distance D between the anode 120 and separator 130 in some aspects is from 5 mm to 20 mm. The gap advantageously provides a high current efficiency for metal hydroxide production, such as a current efficiency >50%, >75%, >90%, or even up to 99%, such as a current efficiency from >50% to 99%, >75% to 99%, or >90% to 99%. The reaction may simultaneously result in acidification of the anolyte. Thus, in some aspects, the pH of the anolyte is increased periodically and/or the anolyte is replenished to maintain an alkaline pH with selective generation of O2 as the primary product in the anode compartment.

In an independent aspect (FIGS. 6B, 6D), the surface 120a of the anode 120 is in direct contact with the separator surface 130a such that there is no gap between the separator 130 and the anode 120. An arrangement with no gap may be preferable when X2 evolution is a desirable co-product. In such arrangements, the anode 120 (e.g., an anode mesh or expanded metal anode mesh) may be compressed against the separator 130. An anode flow plate 620 (see, FIGS. 6D, 7B) may provide an even distribution of pressure across the anode 120 and separator 130. The optimum pressure may depend, at least in part, on the relationship between cell voltage and maximum torque. In some embodiments, the applied pressure to the anode may be from 2 N*m (Newton meters) to 7 N*m.

The catholyte comprises (i) water, (ii) a metal hydroxide comprising Q, where Q is an alkali metal, an alkaline earth metal, or a combination thereof, and (iii) suspended metal ore particles comprising MxOy where M is a metal and x and y are integers. In some embodiments, MxOy comprises a multivalent metal oxide. The integers x and y are determined by the charge on M and, in combination, provide a neutral metal oxide compound. The metal oxide may be sparingly soluble (e.g., a solubility of less than 1 g/1000 mL) in the alkaline environment.

In some implementations, the metal ore has a maximum silicate (e.g., SiO2) content of 5 wt %, such as a silicate content of 0 wt % to 4 wt %, 0 wt % to 3 wt %, or 0 wt % to 2 wt %. In some implementations, the metal ore has a maximum aluminate content of 5 wt %, such as an aluminate content of 0 wt % to 4 wt %, 0 wt % to 3 wt %, or 0 wt % to 2 wt %.

In any of the foregoing or following embodiments, M may be Fe, Mn, Ni, Cr, Co, Zn, or any combination thereof. In certain implementations, M is Fe. When M is Fe, the process may be referred to as the “Chlor-Iron” process. When M is Fe, MxOy comprises Fe2O3, FeO, Fe3O4 (Fe2+Fe3+2O4), or any combination thereof. In some examples, the metal ore is hematite, and MxOy comprises Fe2O3. In any of the foregoing or following embodiments, the suspended metal ore particles may have an average size within a range of from 10 μm to 1 mm, such as from 10 μm to 750 μm, or 10 μm to 500 μm. In any of the foregoing or following embodiments, prior to applying the voltage, the catholyte may comprise from 50 g/L to 500 g/L (or 5 wt % to 50 wt %) of the suspended metal ore particles. In some embodiments, the initial concentration of the suspended ore particles is from 100 g/L to 400 g/L, 100 g/L to 300 g/L, or 150 g/L to 250 g/L, or 10 wt % to 400 wt %, 10 wt % to 30 wt %, or 15 wt % to 25 wt %. In certain examples, the initial suspended ore particle concentration is 20 wt % or 200 g/L.

In any of the foregoing or following embodiments, Q may be Na, K, Li, Cs, Rb, Mg, Ca, or any combination thereof. In some implementations, Q is Na and the metal hydroxide is NaOH. Prior to applying the voltage, the catholyte may comprise from 10 wt % to 50 wt % of the metal hydroxide. In some embodiments, the initial metal hydroxide concentration is 10 wt % to 40 wt %, 20 wt % to 40 wt %, or 25 wt % to 35 wt %. In certain examples, the initial metal hydroxide concentration is 30 wt %. As shown in the cathode reaction above, additional metal hydroxide is produced as the electrochemical cell is operated.

The anolyte comprises water and a halide salt comprising Q and X, where Q and X are as previously defined. In some embodiments, X is Cl or Br. In certain embodiments, X is Cl. In any of the foregoing or following embodiments, Q may be the same in both the catholyte and the anolyte. Thus, if the catholyte comprises NaOH, then the anolyte may comprise NaCl or NaBr. Similarly if the catholyte comprises KOH, then the anolyte may comprise KCl or KBr. In some implementations, Q is Na and the halide salt comprises NaCl. Prior to applying the voltage, the anolyte may comprise from 10 wt % to 50 wt % of the halide salt. In some embodiments, the initial halide salt concentration is 10 wt % to 40 wt %, 20 wt % to 40 wt %, or 25 wt % to 35 wt %. In certain examples, the initial halide salt concentration is 30 wt %. In some implementations, the halide salt is provided by seawater. The seawater may be concentrated (e.g., via evaporation) to provide an initial halide salt concentration of from 10 wt % to 50 wt %, such as an initial halide salt concentration of 10 wt % to 40 wt %, 20 wt % to 40 wt %, or 25 wt % to 35 wt %.

In some implementations, the catholyte comprises water, NaOH, and suspended metal ore particles comprising Fe2O3, and the anolyte comprises water and NaCl. In one non-limiting example, the catholyte comprises 30 wt % NaOH and 20 wt % Fe2O3 in water, and the anolyte comprises 30 wt % NaCl in water.

During operation of the electrochemical cell, a locally alkaline environment (e.g., pH>14) is maintained at the cathode. In some aspects, when halogen production is desired, an acidic environment is maintained at the anode. A bulk average pH≤ 3 favors halogen production in the anolyte by oxidation of X−. In some examples, the bulk average pH is 0-3, such as a bulk average pH of 1-3 or 1.5-2.5, when halogen production is desired. In some embodiments, when oxygen production is desired, a neutral to alkaline environment is maintained at the anode, which shifts the thermodynamic potential making oxygen evolution more favorable than halogen evolution; oxygen is produced by water decomposition. A bulk average pH≥7 favors oxygen production. In some examples, a bulk average pH>14 is preferred for oxygen production. In other examples, an initial bulk average pH of 7-10 is preferred to minimize consumption of metal hydroxide in the anolyte stream. While higher pH values may reduce the equilibrium cell voltage, higher pH values also reduce the rate of metal hydroxide generation. In some aspects, when hypohalite production is desired, a mildly acidic environment, e.g., bulk average 3<pH<7, is maintained at the anode. When bulk average 3<pH<7, a mixture of O2 and X2 is produced and XO− subsequently is formed. The XO− may react with partially reduced metal oxides (e.g., Fe3O4 or Fe(OH)2) to reform X−. The separator facilitates maintenance of the cathode and anode environments. Thus, the desired co-product may be selected by adjusting the anolyte pH.

Conversion and/or dissolution of metal species (e.g., MxOy) occurs locally at the cathode which minimizes or prevents undesirable crossover of M cations through the separator. Additionally, crossover of M cations through the separator may be minimal or non-existent because the reduction of MxOy occurs without producing significant quantities of free M cations. For example, when MxOy is Fe2O3, most or all of the iron in the hematite is directly reduced to Fe; any free unreduced Fe3+ cations are coordinated with hydroxide anions to form Fe(OH)4− anions, which do not pass through the separator. In some aspects, the separator does not permit crossover of M cations.

FIG. 9 illustrates another exemplary electrowinning process 900. Steps 901-904 are the same as those described for FIG. 8, steps 801-804, respectively. In the exemplary embodiment of FIG. 9, the electrowinning process further comprises continuously or periodically removing X2 and/or O2 from the anode compartment (step 905). Thus, steps 904 and 905 are performed simultaneously. In other words, continuous or periodic removal of X2 and/or O2 occurs while applying the voltage. The process of FIG. 9 also comprises periodic removal of at least a portion of the deposited metal M. In some embodiments, from 5-100% of the deposited M is removed from the cathode surface, such as from 10-100%, 25-100%, or 25-75%. In one embodiment, voltage application is ceased (step 906), and at least a portion of the deposited metal M is removed from the cathode compartment (step 907a). Following removal of metal M, the voltage is reapplied (step 904). In an independent embodiment, periodic removal of metal M is performed while applying the voltage (step 907b, performed while step 904 is being performed). The metal M may be periodically removed by any suitable means. In some embodiments, the cathode is removable and deposited metal M is removed from the cathode surface by periodically removing the cathode from the cathode compartment and then scraping, peeling, or rinsing the cathode surface to remove the deposited metal M; in such arrangements, voltage application is ceased (step 906) prior to removing the cathode from the cell. In some implementations, the M is deposited onto the cathode as a dense film. When the metal is dense, it may be removed from the cathode as a single, free-standing film. In certain embodiments, the metal M is magnetic and is deposited onto a surface of the cathode. In some aspects, the deposited magnetic metal is removed as described above. In other aspects, periodically removing at least a portion of the magnetic metal M comprises (i) passing a magnet over a surface of the cathode or over an opposing surface of the cathode, and (ii) removing the magnet from the cathode compartment, whereby the metal M deposited onto the surface of the cathode is transferred to the magnet as the magnet is removed. In such arrangements, metal removal may occur with or without ceasing the voltage application. In certain implementations, the magnet is a magnetic drawbar or a magnetic rod. In some aspects (not shown), if H2 is formed in the cathode compartment, the electrowinning process may further comprise continuously or periodically removing H2 from the cathode compartment at step 905 or in step 907a or 907b.

FIG. 10 is a schematic diagram illustrating one exemplary embodiment of the electrowinning process of FIGS. 8 and 9, wherein the metal ore comprises Fe2O3, the metal hydroxide comprises NaOH, and the halide salt comprises NaCl. An electrochemical cell 100 comprises, in part, a cathode 110, and anode 120, and a separator 130. An aqueous catholyte comprising 30 wt % NaOH and 20 wt % Fe2O3 particles is in contact with the cathode 110, and an aqueous anolyte comprising 30 wt % NaCl is in contact with the anode 120. Upon application of a voltage across the electrochemical cell 100, the iron in Fe2O3 is reduced and deposited as Fe metal on a surface 110a of the cathode 110. Concurrently, chloride ions are oxidized at the anode 120 to form Cl2 and/or O2 gas. In some aspects, a removable cathode 110 may periodically be removed from the electrochemical cell 100 and the deposited metal M removed from the cathode surface 110a by any suitable means (e.g., manual scraping, peeling, or rinsing).

In some implementations, a magnet 180 is passed over an opposing surface 110b of the cathode to remove Fe from the surface 110 a. With reference to FIGS. 2C, 3A, 3B, and 10, as the magnet 180 passes over the cathode surface 110b, the deposited magnetic metal M on the cathode surface 110a moves along the surface 110a as the magnet moves along the opposing surface 110b. As the magnet 180 is withdrawn from the electrochemical cell 100, at least a portion of the metal M is transferred from the cathode surface 110a to a surface of the magnet 180 and removed from the electrochemical cell 100. With reference to FIG. 3C, as the magnet 180 passes over the cathode surface 110a, deposited metal M on the cathode surface 110a is transferred to the magnet 180. In any of the foregoing or following embodiments, the magnet 180 may be manually passed over the cathode surface, or the process may be automated. In some embodiments, from 5-100% of the deposited M is removed from the cathode surface, such as from 10-100%, 25-100%, or 25-75%. When the magnet 180 is immersed in the catholyte (e.g., as shown in FIG. 3C), metal M fines that are not deposited onto the cathode surface 110a and are free in the catholyte 140 also may be removed by the magnet 180. Advantageously, when the metal ore is non-magnetic (e.g., hematite), the metal ore particles remain in the catholyte 140.

Cl2 and/or O2 may be continuously or periodically removed from the anode compartment 120 by removing accumulated Cl2 and/or O2 gas from the gas envelope 170 via the outlet 172 (FIG. 2C). Advantageously, when the gas envelope 170 is configured to extend only over the anode compartment 120, the metal M may be removed from the cathode compartment 110 without opening the anode compartment 120.

As the electrowinning process is performed, the concentration of metal ore particles in the catholyte and the concentration of halide salt in the anolyte decreases. Thus, in some embodiments, the catholyte and anolyte are periodically replenished to restore desired concentrations of metal ore particles and/or halide salt. Additionally, in certain implementations, such as when the separator is a porous composite, some hydroxide anions may cross through the separator and into the anolyte, thereby reducing the hydroxide concentration in the catholyte and increasing a pH of the anolyte. In such implementations, the method may further include replenishing metal hydroxide in the catholyte and/or adding dilute HX (HCl, HBr) to the anolyte to maintain a desirable pH.

FIG. 11 illustrates an exemplary electrowinning process 1100 in which the catholyte and anolyte are periodically replenished. Steps 1101-1104 are the same as those described for FIG. 8, steps 801-804, respectively. In the exemplary embodiment of FIG. 11, the electrowinning process further comprises (i) periodically adding a quantity of the metal ore particles and/or a quantity of metal hydroxide to the catholyte; or (ii) periodically adding a quantity of the halide salt and/or dilute HX (e.g., 0.1-6 M HX) to the anolyte (to maintain a slightly acid pH at the anode); or (iii) both (i) and (ii). In some embodiments, the electrowinning process is a batch process and voltage application ceases (step 1105) while the catholyte and/or anolyte is replenished (step 1106a). Voltage application (step 1104) may resume after replenishment (step 1106a) is complete. In an independent embodiment, the electrowinning process is a continuous process and catholyte and/or anolyte replenishment (step 1106b) is performed while the voltage continues to be applied (step 1104). In such implementations, the electrochemical cell may be a flow cell. In some embodiments, adding the quantity of the metal ore particles comprising adding a suspension comprising the quantity of the metal ore particles to the catholyte. In some implementations, adding a quantity of the halide salt comprises adding a solution or a suspension comprising the quantity of the halide salt to the anolyte. In some implementations, periodically replenishing the catholyte with additional metal ore particles and periodically replenishing the anolyte with additional halide salt allows the process to continue without completely replacing the catholyte and anolyte.

When the metal ore particle concentration decreases, the metal hydroxide concentration in the catholyte concomitantly increases, particularly when the separator is a cation-selective membrane that does not allow passage of hydroxide anions. Replenishing the catholyte by adding a suspension of metal ore particles also dilutes the metal hydroxide concentration. When the separator is a cation-selective membrane, the dilution may restore a desired concentration of the metal hydroxide. However, if the separator allows crossover of hydroxide and the metal hydroxide concentration is insufficient after dilution, a quantity of metal hydroxide may be added as a solid, a suspension, or a solution, and may be added together with or separately from the metal ore particles to restore the desired hydroxide concentration (steps 906a, 906b). If dilute HX is added to the anolyte to compensate for any hydroxide crossover, the dilute HX may be added together with or separately from the halide salt (steps 1106a, 1106b). It is understood by a person of ordinary skill in the art that, in any of the foregoing or following embodiments, the process may further include removing a volume of the spent catholyte and/or anolyte from the electrochemical cell before replenishing the electrolytes so that the cell does not overflow.

FIG. 12 illustrates an exemplary electrowinning process including both product removal and electrolyte replenishment. Steps 1201-1204 are the same as those described for FIG. 8, steps 801-804, respectively. Steps 1205, 1206, 1207a, and 1207b are the same as those described for FIG. 9, steps 905, 906, 907a, and 907b, respectively. Steps 1206, 1208a and 1208b are the same as those described for FIG. 11, steps 1105, 1106a and 1106b, respectively. In some aspects (not shown), if H2 is formed in the cathode compartment, the electrowinning process may further comprise continuously or periodically removing H2 from the cathode compartment at step 1205 or in step 1207a or 1207b.

The steps of product removal and electrolyte replenishment may be performed at the same frequency or with different frequencies. For example, removal of metal M may occur more frequently than electrolyte replenishment, or electrolyte replenishment may occur more frequently than removal of metal M. Addition of metal hydroxide and/or dilute HX may occur at the same frequency or less frequently than addition of metal ore particles and/or halide salt.

In one embodiment, step 1206 is followed by step 1207a and then by step 1208a, after which the process returns to step 1204. In an independent embodiment, step 1206 is followed by step 1208a and then by step 1207a, after which the process returns to step 1204. In another independent embodiment, only one of steps 1207a and 1208a is performed before returning to step 1204. In yet another independent embodiment, steps 1207b and 1208b are performed substantially consecutively in any order while the voltage continues to be applied. In still another embodiment, only one of steps 1207b and 1208b is performed and an interval of time passes before the other of steps 1207b and 1208b is performed; the voltage is applied throughout the process. When multiple cycles of steps 1204-1208 are performed, the selection and sequence of steps 1206/1207a/1208a or 1207b/1208b may vary with each cycle.

In any of the foregoing or following embodiments, providing the electrochemical cell (steps 801, 901, 1101, 1201) may further comprise providing a cell stack. In one embodiment (e.g., as illustrated in FIG. 4), the cell stack comprises a number n of the electrochemical cells, a cathode electrical connector connecting cathodes of each of the electrochemical cells in parallel, an anode electrical connector connecting anodes of each of the electrochemical cells in parallel, and a voltage source electrically connected to the cathode electrical connector and the anode electrical connector. In an independent embodiment (e.g., as illustrated in FIG. 5), the cell stack comprises a number n of the electrochemical cells, a number n−1 of conductive bipolar plates wherein one conductive bipolar plate is positioned between each adjacent pair of electrochemical cells, a cathode electrical connector connected to a cathode of a first electrochemical cell in the series, an anode electrical connector connected to an anode of a last electrochemical cell in the series, and a voltage source electrically connected to the cathode electrical connector and the anode electrical connector. In any of the foregoing or following embodiments where the process is performed with a cell stack, providing the catholyte within the cathode compartment (steps 802, 902, 1102, 1202) may further comprise providing the catholyte within each cathode compartment of the cell stack. In any of the foregoing or following embodiments where the process is performed with a cell stack, providing the anolyte within the anode compartment (steps 803, 903, 1103, 1203) may further comprise providing the anolyte within each anode compartment of the cell stack. As previously discussed, the catholyte and anolyte in each cathode compartment and anode compartment, respectively, may have the same composition or different compositions. In any of the foregoing or following embodiments where the process is performed with a cell stack, applying a voltage across the electrochemical cell may further comprise applying the voltage across the cell stack to effect reduction of the metal ore (MxOy) particles in each cathode compartment to provide the metal M and formation of Cl2 gas in each anode compartment. If the catholyte and anolyte compositions vary from cell to cell within the stack, the deposited M may be a different metal in one or more the cells.

In any of the foregoing or following embodiments, the electrochemical cell or cell stack may be operated at a temperature of from 25° C. to 150° C., such as a temperature in a range having endpoints selected from 25° C., 30° C., 35° C., 40° C., 50° C., 60° C., 70° C., 80° C., 90° C., 100° C., 110° C., 120° C., 130° C., 140° C., or 150° C., wherein the range is inclusive of the endpoints. In some embodiments, the electrochemical cell or cell stack is operated of from 50° C. to 70° C., 70° C. to 90° C., 90° C. to 120° C., or 100° C. to 150° C. An operating temperature >60° C. may be utilized when X is Br and Br2 is a desired co-product. The concentrated catholyte and anolyte enable the electrochemical cell to operate at temperatures above the boiling point of water. However, the temperature may be maintained at temperatures lower than 150° C. (e.g., 25° C. to 100° C.) to prolong the lifetime of various cell components (e.g., the anode and/or separator) that may fail at more elevated temperatures.

In any of the foregoing or following embodiments, the voltage applied may be from 2 V to 5 V, such as 2 V to 4 V or 2.5 V to 4 V (3500 to 5500 kWh tonne iron), for a single cell. In any of the foregoing or following embodiments, the electrochemical cell or cell stack may be operated at a current density of from 20 mA cm−2 to 500 A cm−2, such as a current density range having endpoints selected from 20 mA cm−2, 50 mA cm−2, 100 mA cm−2, 250 mA cm−2, 500 mA cm−2, 1 A cm−2, 5 A cm−2, 10 A cm−2, 25 A cm−2, 50 A cm−2, 100 A cm−2, 200 A cm−2, 300 A cm−2, 400 A cm−2, or 500 A cm−2, wherein the range is inclusive of the endpoints. In some embodiments, the current density is 20 mA cm−2 to 400 A cm−2, 20 mA cm−2 to 300 A cm−2, 50 mA cm−2 to 300 A cm−2, or 50 mA cm−2 to 200 A cm−2. The current density may depend, at least in part, on the applied voltage, the electrochemical cell temperature, the concentrations of the metal ore particles comprising MxOy and the halide salt, or any combination thereof. When the device comprises a stack of cells connected in parallel (e.g., as shown in FIG. 4), the stack current is calculated as the single-cell current×the number of cells.

In any of the foregoing or following embodiments, hydrogen co-production is increased as the current density increases beyond 300 mA cm−2 and/or when the metal oxide particles are larger than 0.01 mm in diameter and lack internal surface area available for penetration with electrolyte.

In any of the foregoing or following embodiments, the current efficiency toward metal ore reduction at the cathode (over side reactions, such as hydrogen evolution) may range from 50% to 95%, such as from 60% to 95%, 70% to 95%, 80% to 95%, or even 85% to 95%.

In any of the foregoing embodiments, the process may be a batch process, a continuous process, or a semi-continuous. In a continuous or semi-continuous process, product removal and/or reactant replenishment may be performed while continuing to apply the voltage across the electrochemical cell. For stacks of cells that are electrically connected in parallel as opposed to series, metal can be harvested from a current collector while the remaining cells in the stack remain in operation. This maximizes the capacity factor for the cells and exerts a negligible influence on the average stack current, provided there are a significant fraction of cells left in operation. In some aspects, use of a magnet, such as a magnetic drawbar, to collect the deposited metal M, may facilitate continuous operation of the electrochemical cell or stack. In a continuous or semi-continuous process, pumps that circulate or recirculate the catholyte may become fouled by the suspension of metal ore particles and eventually fail. In some embodiments, the process is a batch process. In the batch process, the voltage may be applied for a period of time, after which voltage application ceases, the metal M is removed from the cathode compartment (Cl2 is removed continuously from the anode compartment during cell operation). The catholyte and/or anolyte may be replenished, and cell operation is resumed.

Embodiments of the disclosed processes exhibit several advantages over prior electrowinning processes. For example, the reactants may include inexpensive materials, such as iron oxide ores (e.g., hematite) and salt water brines. Although an initial quantity of metal hydroxide is needed, the net reaction (MxOy(s)+yH2O+2yX−→2yOH−+xM(s)+yX2(g)) produces hydroxide during the process. Advantageously, the process results in two or more commodity products. In one aspect, the products are the metal M and X2 gas—produced at large scale (at least 10 Mt/y within the U.S.). In an independent aspect, the products are the metal M and O2 gas. In another independent aspect, the products are the metal M and the hypohalite XO−. In yet another independent aspect, the products may further include H2 or, when H2 and X2 are co-produced, the H2 and X2 may be combined to spontaneously form HX.

In some implementations, when M is Fe and the metal ore does not include significant amounts of silicates (e.g., the metal ore includes no more than 5 wt % silicates), the electrowinning process may have a Faradaic efficiency toward iron (ηFe) of at least 85% at current densities from 100 mA cm−2 to 300 mA cm−2. In certain aspects at current densities ≤200 mA cm−2, ηFe is >90%, such as >94%. In some implementations, the Faradaic efficiency toward Cl2 (ηCl2) is at least 80%, at least 85%, or at least 90%.

Reduced metal M is recovered as a solid from the cathode compartment. In some aspects, X2 and/or O2 gas is recovered from the anolyte compartment. The co-produced hydroxide is dissolved in the catholyte. Separating these products is straightforward since they are present as different states of matter.

In some embodiments, the recovered metal M has a purity of at least 90 wt %, at least 95 wt %, at least 96 wt %, at least 97 wt %, or at least 98 wt %. In some aspects, the reduced metal comprises particles having an average diameter of from 5 μm to 100 μm, such as from 10 μm to 50 μm. In other aspects, the metal is deposited as a film, such as a free-standing film. The morphology may be controlled, at least in part, by the applied current density. In one non-limiting example, a current density of 100 mA cm−2 produced films largely free of pinholes, whereas deposition at 300 mA cm−2 provided a product comprising multiple large pinholes. When M is Fe, the iron produced by aspects of the disclosed electrochemical cell and process has a morphology and purity suitable for use as feedstock in electric arc furnaces for steel manufacturing or production of monolithic iron electrodes, among other uses. The halogen gas may be used the production of halide salt salts, or preparing halogen-containing compounds, as well as other utilities, such as use in bleaching processes, disinfecting processes, and water cooling systems, among others. When O2 of high purity (>99%) is the product of the anode, it can be vented to the atmosphere, used to enhance combustion processes such as roasting of ores, or reacted with dissolved transition metals (e.g. Fe2+) to produce high purity metal oxides. When H2 and X2 are co-produced and combined to form HX, the acid may be used for metal ore processing, among other uses. Advantageously, acids facilitate selective dissolution of Fe cations from impure sources of iron oxide, which can then be reacted with metal hydroxide and/or O2 from the air to form higher purity Fe2O3 sediments which are highly reactive for direct reduction of iron oxide to iron metal.

The hydroxide produced during the process is a third commodity product. The spent catholyte comprises a solution of metal hydroxide and water, and may further comprise residual metal ore (MxOy) particles. While the hydroxide solution may comprise impurities since the spent catholyte may include residual amounts of metal oxide or metal, the low-purity hydroxide nonetheless has multiple uses. One use is purification of metal ore, such as leaching out impurities. For example, if a metal ore comprises silicates (e.g., SiO2) and/or aluminates, the metal hydroxide solution may be used to leach at least some of the silicates and/or aluminates from the metal ore. In some implementations, metal ore intended for use in the disclosed electrowinning process may be pre-leached with metal hydroxide from the spent catholyte prior to introducing the metal ore into the electrochemical cell. In another aspect, the metal hydroxide solution may be used for carbon capture by capturing and mineralizing CO2, such as by stoichiometrically capturing and mineralizing CO2 from the air or ocean. In some implementations, the spent catholyte comprising metal hydroxide is used (e.g., for leaching or carbon capture) without further purification. In certain implementations, the spent catholyte may be filtered before use to remove any residual metal and/or metal particles, and/or otherwise purified.

The direct use of electricity as a reductant circumvents the need for shipping and storing intermediate reducing agents such as hydrogen and thus provides a pathway to lower plant costs. The operating temperature (room temperature to 150° C.) may be lower than in other processes, which allows use of less expensive materials for the electrodes, separators, and other components than processes operating at more extreme temperatures. The low operating temperatures are compatible with intermittent zero-emissions electricity (e.g., solar energy, wind energy). Embodiments of the disclosed process also eliminate the need to dissolve metal salts, the need to perform an initial reduction process (e.g., Fe3+ to Fe2+), and the need to transport and store large quantities of HCl and metal hydroxides. In traditional processes requiring metal salt dissolution, electrowinning of multivalent metals (including Fe, Ni, or Cr) can lead to parasitic currents between the anode and cathode and fouling of the separator by salt crossover is common. Parasitic currents and separator fouling may be reduced or eliminated by embodiments of the disclosed electrochemical cell and method.

IV. EXAMPLES Example 1 Cell Preparation and Electrolysis Protocol

Anode preparation: An ink for preparing a ruthenium oxide coating was prepared by dissolving RuCl3 (anhydrous) powder in 20% w/w hydrochloric acid and applying gentle heat until all the liquid had been removed. The powder that remained was resuspended in isopropyl alcohol to a concentration of approximately 2 mg/mL. This liquid (hereafter “ink”) was used to applying ruthenium, Ru, to the electrode surface. Titanium was polished with micron-grade alumina polishing paper, soaked in 20% w/w hydrochloric acid for 1 min and rinsed with deionized water. The substrate was then massed. Titanium was repeatedly brushed with the ink and dried on a hotplate set to 130° C. and then annealed in an open-air furnace set to 500° C. in 10-15 minute increments with the loading of the Ru checked by changes in mass until the added mass of the coating was 4 mg.

Cathode preparation: Iron was plated onto low carbon steel substrates that polished with micron-grade alumina polishing paper and then massed prior to use as a deposition substrate.

Cell Assembly: A cell 1300, as illustrated in FIG. 13, was assembled using ¼″ (0.64 cm) stainless steel bolts and nylon washers. The cell included graphite blocks 1310 at each end, an anode 1320, a cathode 1330, and a series of Viton™ (DuPont, Wilmington, DE) synthetic rubber and fluoropolymer elastomer gaskets 1340 and polyether ether ketone (PEEK) spacers 1350. In the exemplary cell, the graphite blocks 1310 were 8 cm square and 1.27 cm thick. The gaskets 1340 were 8 cm square and 0.05 cm thick. The spacers 1350 were 8 cm square and 0.635 cm thick. Silicone resistive heating mats were fixed to either side of the cell and the temperature was monitored with an embedded thermocouple in the graphite plate. A Nafion™ 117 (DuPont) fluorinated membrane (not shown) stored in deionized water was placed between the two PEEK spacers 1350 to divide the cell 1300 into an anode and cathode compartment.

Electrolysis: The cell was heated in the dry state to a temperature of 80° C. at which point the anolyte and catholyte were added to the chamber. All solutions were prepared with deionized water having a resistivity >18 MΩ. The anolyte consisted of 20% w/w NaCl solution in deionized water. The catholyte consisted of a 30% w/w NaOH solution in deionized water. 2.5 mL of solution was mixed with 0.3 g of Fe2O3 powder prior to addition to the cathode chamber. Current and voltage supplied to the cell was controlled with a BioLogic SP300 Potentiostat (Seyssinet-pariset, France). Bulk electrolysis was carried out at various applied current densities for 30 or 60 minutes with the faradaic efficiency of Fe plating estimated via changes in mass at the cathode after the electrode had been rinsed thoroughly in deionized water and briefly dried at 100° C.

Half-Cell Analysis

Kinetic overpotentials for the Fe plating reaction and chlorine evolving reaction were measured at a graphite cathode and Ti/RuO2 anode (Example 1), respectively, to obtain information about the practical minimum cell voltage in the absence of resistances through the electrolyte and membrane but accounting for kinetic overpotentials. The catholyte consisted of 30% w/w NaOH+Fe2O3 in deionized water (Example 1), and the anolyte consisted of 20% w/w NaOH in deionized water. The current-voltage behavior of a graphite cathode in 30% w/w NaOH in the absence of Fe2O3 has been included for comparison. The results are shown in FIG. 14.

The data support the predicted practical range of operating voltages. For a cell resistance between 3-7 Ω/cm2 (inclusive of electrolyte resistance and leads) operating at currents between 0.1 and 0.3 A/cm2 the predicted cell voltage would be between 2-3.8 V. This was calculated by adding the corresponding minimum and maximum ohmic overpotentials to the measured practical minimum cell voltage. Implicit is an assumption that the kinetic overpotentials will contribute negligibly compared to ohmic overpotentials and that mass transfer overpotentials will be insignificant in the concentrated electrolytes used.

Full Cell Electrolysis

The cell of FIG. 13 was assembled and tested at constant current at a temperature of 80° C. The measured cell resistance was 3.6 ohms in the absence of evolved gas bubbles (which are known to contribute substantially to the cell resistance under operation). The approximate geometric area of the cell was 2 cm2 defined by the area of the two metallic electrodes. A constant current of 100 mA (equivalent to a nominal current density of 50 mA/cm2 was applied) for 30 min at a time-averaged cell voltage of 2.9 V (FIG. 15A). The formation of iron-sponge at the cathode was evident (FIG. 15B) and gases were continuously evolved at both the anode and cathode but not yet collected for quantification. The selectivity towards Fe production was 51% based on the measured change in mass at the cathode relative to the expected change in mass for a 100% conversion of all charge passed toward the three-electron reduction of Fe2O3 to Fe metal.

Example 2

Materials: α-Fe2O3 (98% purity, metals basis), NaOH (Extra Pure, 50.0 wt % aq), Al2O3 (99.7%, <0.04% SiO2), and anhydrous RuCl3 (47.7% min), were purchased from Thermo Scientific, whereas HCl (ACS Grade), and NaCl (ACS Grade) were purchased from Fisher Scientific. SiO2 was purchased from U.S. Nano. Nafion™ 115 and GI-N417 membranes were purchased from Fuel Cell Store, a Zirfon® UTP-500 membrane was provided by Agfa, and a Nafion™ N2030 membrane was provided by Chemours. Ti foil (99.99%) was purchased from Futiantian Technology Co. and Cu foil (99.99%) from MTI Corporation. A representative commercial-grade iron oxide powder was purchased from Alpha Chemical. KBr was sourced from J.T. Baker and Methyl Orange (97%) from Aldrich. Iron (III) nitrate nonahydrate Fe(NO3)3·9H2O (98+%) was purchased from Strem Chemicals (Newburyport, MA). All solutions were prepared using water from a Millipore® deionized water system having a resistivity >18.1 MΩ.

Anode preparation: Ti-supported RuOx served as Cl2 evolving anodes and were prepared following the method of Trasatti (J. Electroanal. Chem. Interfacial Electrochem. 1971, 29(2):A1-A5). Briefly, 80 mg of RuCl3 was dissolved in 20 mL of 20 wt % HCl(aq) and then heated to evaporate excess liquid, yielding a dark residue, free of visible particles. This residue was resuspended in 5 mL of isopropyl alcohol to form a Ru(III) ink and stored in the dark at room temperature until use. Ti foils were soaked in 20 wt % HCl(aq) for 10 min and then repeatedly brush-coated with the Ru(III) ink, dried on a hot plate at 150° C., and suspended in stagnant air within an upright furnace at 400° C. for 10 min. The coating, drying, and oxidation process was repeated until the catalyst loading was ˜2 mg/cm2.

Cell Assembly: A cell 1300, as illustrated in FIG. 13, was assembled using ¼″ (0.64 cm) stainless steel bolts and nylon washers. The cell included graphite blocks 1310 at each end, an anode 1320, a cathode 1330, and a series of Viton™ (DuPont) synthetic rubber and fluoropolymer elastomer gaskets 1340 and polyether ether ketone (PEEK) spacers 1350. In the exemplary cell, the graphite blocks 1310 were 8 cm square and 1.27 cm thick. The gaskets 1340 were 8 cm square and 0.05 cm thick. The spacers 1350 were 8 cm square and 0.635 cm thick. Silicone resistive heating mats were fixed to either side of the cell and the temperature was monitored with an embedded thermocouple in the graphite plate. A Nafion™ 117 (DuPont) fluorinated membrane (not shown) stored in deionized water was placed between the two PEEK spacers 1350 to divide the cell 1300 into an anode and cathode compartment.

Electrolysis: The cell was heated in the dry state to a temperature of 80° C. at which point the anolyte and catholyte were added to the chamber. All solutions were prepared with deionized water having a resistivity >18 MΩ. The anolyte consisted of 20% w/w NaCl solution in deionized water. The catholyte consisted of a 30% w/w NaOH solution in deionized water. 2.5 mL of solution was mixed with 0.3 g of Fe2O3 powder prior to addition to the cathode chamber. Current and voltage supplied to the cell was controlled with a BioLogic SP300 Potentiostat (Seyssinet-pariset, France). Bulk electrolysis was carried out at various applied current densities for 30 or 60 minutes with the faradaic efficiency of Fe plating estimated via changes in mass at the cathode after the electrode had been rinsed thoroughly in deionized water and briefly dried at 100° C.

Example 3

The materials and method of anode preparation were the same as in Example 2 but with a modified design for the electrochemical cell and operating procedure as detailed below.

Electrochemical cell and characterization: The electrochemical cell was prepared from two graphite plates supporting a polished Cu foil cathode, a Ti/RuOx anode, neoprene rubber gaskets, polyether ether ketone (PEEK) plates, and a cation-exchange membrane. The anode and cathode compartments were defined by 6 mm thick polyether ether ketone (PEEK) spacers and were separated by a cation-selective membrane. A bilayer membrane (Nafion™ N2050 membrane, Chemours, Wilmington, DE), designed to prevent hydroxide crossover in commercial chlor-alkali cells was used to separate the compartments of the flow cell. All cells were externally heated to 80° C. through a silicone heating mat (Omega Engineering) fixed to graphite endplates in contact with the current collectors. A thermocouple was inserted into the graphite plate to measure the cell temperature and give feedback to a simple controller providing power to the heating pads.

Current was supplied to the cell from a programmable DC power supply and voltage was recorded using a Biologic SP200 potentiostat. The standard catholyte was an alkaline suspension of 25 wt % α-Fe2O3 mixed with 7.5 mol/kg NaOH(aq) and the anolyte was 5.7 mol/kg NaCl(aq) adjusted to pH=2 using 1 M HCl(aq). The effects of impurities were measured with a 25 wt % suspension of SiO2/Al2O3/Fe2O3 in 7.5 mol/kg NaOH(aq), where the mass fraction of SiO2 or Al2O3 was expressed relative to the mass of the solids prior to addition to the basic catholyte.

Cathode conditions: A batch cell separated by a Nafion™ 115 (Chemours) membrane was used to determine optimal current density conditions for Fe plating. At J<100 mA cm−2, a BioLogic SP300 potentiostat (BioLogic, Seyssinet-pariset, France) supplied the current and recorded voltage. For J>100 mA cm−2, a 30V-10 A benchtop power supply provided the provided current and a Biologic SP300 recorded voltage. The temperature of the cell was monitored with a negative temperature coefficient thermistor and controlled using a voltage relay connected to silicone heating mats integrated with the cell. The plating efficiency was determined from the weight of the washed and dried cathode post-plating (Equation 1), as the composition of Fe by mass was consistently in excess of 95%.

η Fe = m Fe ⁢ zF QM Fe ( 1 )

where mFe is the measured change in mass, z is the number of electrons (assumed 3 for purified Fe2O3), F is Faraday's constant, MFe is the molar mass of iron, and Q is the charged passed during electrolysis.

Impurity testing: The source of Fe used for the catholyte suspension impacted the performance of the cell. A high purity source of Fe2O3 (Thermo Scientific) was used as the baseline and led to a substantially increased faradaic efficiency towards Fe in comparison to a lower purity red iron oxide powder (Alpha Chemicals). The effect of metal oxide impurities on Fe electrolysis was studied using a factorial experiment with a single replicate and a randomized run order that screened for primary and cross effects of SiO2, Al2O3, and current density J. The maximum concentration of metal oxide was selected based on the concentrations present in the “commercial grade” ore (8 wt % SiO2 and 4 wt % Al2O3), which was found to lead to low (˜15%) current efficiency towards Fe. Further investigation of SiO2 contamination was conducted after it was identified as a critical parameter influencing ηFe. Despite the direct, inverse relationship between the SiO2 content in the electrolyte and the plating efficiency of Fe, Si was not detected in the films using an electron probe microanalyzer (EPMA).

Symmetric cell for anode evaluation: A mirrored cell with NaCl acid solution on both sides of the Nafion™ membrane was used to evaluate the dependence of ηCl2 on J. For safety, Cl2(g) leaving in the anode effluent was immediately quenched in >10 wt % NaOH. After the anode gas production cleared the effluent tube, the tube was rinsed with DI water and then submerged in the methyl orange solution to quantify Cl2(g) production via endpoint detection.

Quantifying chlorine production: Chlorine produced in a flow cell separated by an N2030 membrane (Chemours) was quantified using endpoint detection via oxidation of a methyl orange indicator solution (Wintrich et al., ChemElectroChem 2019, 6(12):3108-3112). A stock solution was prepared by dissolving 0.050 g of methyl orange and 0.25 g of KBr in 0.050 L of water and then mixing with 0.010 L of ethylene glycol. A stock solution was prepared by diluting 1:3 (v/v) using 0.100 L of 0.50 M H2SO4(aq) and 1.0 mL of 3.0 wt % H2O2(aq). The indicator solution was prepared by diluting the stock solution to 1 L with deionized H2O and Cl2 was quantified using endpoint detection in 0.100 L aliquots of the indicator solution. The transition during titration with Cl2 gas was rapid, occurring in less than a second. During a typical electrolysis experiment at 0.100 A/cm−2, the endpoint was reached in ˜20 seconds such that the faradaic efficiency estimate represents an average efficiency during 20 seconds of electrolysis. Cl2(g) produced in the cells was immediately reacted with titrant or quenched in a chilled, alkaline bath prior to disposal.

Results

A batch-reactor prototype similar to the cell of FIG. 13 was prepared with a cation-selective Nafion™ 115 membrane as the separator. The iron source was α-Fe2O3 aggregate particles having a diameter in micrometers and composed of primary nanoparticles (FIGS. 16A, 16B), that were suspended in a strongly basic electrolyte of 7.5 mol/kg NaOH, pH>14 The metallic cathode product was collected as a solid metal film attached to the current collector (FIGS. 17A, 17B). The anode chamber contained an acidified brine (5.7 mol/kg NaCl, pH 2). A thin film of RuOx, prepared via thermal decomposition of RuCl3 at 400° C. on a Ti current collector, served as the anode (FIGS. 18A, 18B). The cell was externally heated to 80° C. and operated at 100-300 mA cm−2.

The design replaces H2 evolution at the cathode with the direct reduction of metal-oxide particles to metal (Equations 2-4) and simultaneously produces Fe(s), NaOH(aq), and Cl2(g), which are naturally phase separated within the reactor for collection. Fe is reduced from Fe2O3 and NaOH is co-produced at the cathode, while NaCl is oxidized to Cl2 at the anode.

A general summary of the process is illustrated in FIG. 2A. FIG. 19A shows the polarization (current-voltage) behavior of the chlor-iron process after two separate 30 min electrolysis experiments at 100 mA cm−2 and 300 mA cm−2. The onset of the Faradaic efficiency towards Fe, ηFe was evaluated based on the mass of Fe collected from the current collector (equation 1).

The product purity and phase composition Fe films were analyzed using X-ray diffraction (XRD), electron probe microanalysis (EPMA), and energy-dispersive X-ray analysis (EDX). The Fe content in the film as measured by EPMA, 96.8(6)%, and (EDX), 98.3(7)%, was consistent with films composed primarily of metallic Fe with a small amount of oxide present either on the film surface (FIG. 20) or as inclusions within the film. XRD patterns (FIG. 19B) of the reactant powder (upper tracing) were compared to a crushed product film after the top layer was scraped from the surface (lower tracing) showing only the presence of α-Fe2O3 in the reactant while the metallic product exhibited a single peak assigned to α-Fe(110).

The morphology of reduced Fe was dependent on J with films composed of Fe particles tens of micrometers in diameter (FIG. 21C). Films deposited at 100 mA cm−2 were largely free of pinholes with limited branching within the film microstructure. Secondary particles of Fe within the film structure were generally tens of micrometers in diameter (FIGS. 19C, 21A, 21C). The morphology of films produced at 300 mA cm−2 contained a surface base layer with multiple 30 μm diameter pinholes (FIGS. 19D, 21B, 21D) that extended to the current collector surface.

The hydrogen evolution reaction and oxygen evolution reaction compete with the cathodic and anodic reactions in the chlor-iron cell, respectively, and so the selectivity towards the desired Fe and Cl2 products at the RuOx surface as a function of J was evaluated (Adiga et al., Mater. Today Energy 2022, 28:101087). As shown in FIG. 22A, ηFe>85% was sustained for current densities 100-300 mA cm−2. The selectivity towards iron plating, ηFe, was >94% for J≤200 mA cm−2 and ηFe decreased to 87% at 300 mA cm−2. Further increases to J produced excess H2(g) and the resulting gas bubbles prevented reliable cell-voltage measurements within the batch cell, so ηFe was not quantified at J>300 mA cm−2. For the batch cell, the selectivity towards Cl2(g) evolution was evaluated in a cell containing 5.7 mol/kg NaCl (pH=2) in both the anolyte and catholyte, with ηCl2=91(3)% at J=125 mA cm−2 (FIG. 22B).

To evaluate the impact of common contaminants, SiO2 and Al2O3 particles were blended with purified α-Fe2O3 as feedstocks for the chlor-iron cell. At J=100 mA cm−2, ηFe was inversely proportional to the quantity of SiO2 fed to the cell (FIG. 22C), although the polarization behavior was largely unaffected by SiO2. At an increased J of 300 mA cm−2, 8 wt % SiO2 in the catholyte led to an ηFe<37% during electrolysis (Table 2). Each condition was measured with one replicate, with ηFe measured after 30 minutes of continuous electrolysis. The addition of Al2O3 at 4 wt % did not significantly impact ηFe at either J.

TABLE 2 Measured ηFe as a function of intentionally added SiO2 and Al2O3 impurities to reactant Fe2O3. SiO2 wt % Al2O3 wt % ηFe (%) ηFe vs metal oxide concentration at 100 mA cm−2 0 0 94, 95 0 4 95, 95 8 0 58, 57 8 4 57, 54 ηFe vs metal oxide concentration at 300 mA cm−2 0 0 87, 87 0 4 88, 93 8 0 36, 37 8 4 39, 41

An electrochemical flow cell that maintained constant electrolyte volumes and concentrations was prepared for longer duration evaluation of the chlor-iron process and also enabled continuous collection of Cl2(g) for measurement of ηCl2 prior to quenching. The direction of flow within the cell was aligned vertically with respect to gravity during operation with electrolyte pumped in through the base of the cell. Expelled electrolyte due to gas evolution in the cathode chamber was collected and stored, while Cl2(g) was quenched in either a methyl-orange indicator solution used to measure ηCl2 via endpoint detection (Wintrich et al., ChemElectroChem 2019, 6:3108-3112), or an alkaline solution that rapidly converted Cl2(g) to soluble HCIO/CIO−. Pre-heated catholyte and anolyte were slowly pumped into the flow cell which was maintained at a fixed temperature using external silicone heating pads. Gaseous effluent from the anode compartment was quenched in a large beaker containing aqueous NaOH until bubbling reached a steady state. The effluent was then passed to a 1 L beaker containing a stirred methyl orange indicator solution, which rapidly bleached upon reaching a fixed endpoint. The time required to reach the endpoint was used to calculate a molar flux of chlorine which was compared to the applied J to calculate a Faradaic efficiency for the cell.

A bilayer membrane designed for the chlor-alkali process (N2050, Chemours) maintained an acidic pH environment at the anolyte, whereas the anolyte pH rapidly increased to >10 in cells prepared with Nafion™ 115 membranes or porous diaphragms. Both the Nafion™ N2050 and Nafion™ 115 membranes led to cells with similar polarization responses, whereas a porous diaphragm separator led to substantially lower cell voltages (FIGS. 23A-23B). In all cases, linear ohmic losses dominated the polarization response. The cell voltage was consistent at ˜3.2 V throughout a 4 h constant current experiment at 100 mA cm−2 whereas the measured ηCl2 ranged from 82-105% during the first 20 min of reaction before stabilizing at 95(3)% from hours 1-4 and the average ηFe was 84%. The polarization behavior of the cell after testing was similar to the polarization behavior measured after 30 min (FIG. 24). The Fe films produced in multi-hour reactions were easily detached from the current collector surface as a free-standing film (FIGS. 25C-25E). SEM and optical microscopy revealed films mostly consisting of dense Fe metal particles, having diameters ˜10-50 μm with a sparse population of voids having diameters <10 μm (FIG. 25D). The thickness of the film was mostly uniform, except for small clumps that formed within the slurry.

Industrial electrowinning processes rely on the repeated collection of metal from the production cell. A bifacial cathode was used to test the feasibility of collecting Fe product without cell disassembly. To isolate the performance of the cathode and eliminate the influence of pH gradients between anolyte and catholyte, which gradually developed in the absence of controlled flow or recirculating electrolyte in the small area research cell, the supporting electrolyte supplied to the anode and cathode compartments was 7.5 mol/kg NaOH such that the anodes evolved only O2 gas. Consistent with prior experiments using a semi-batch cell and flow cell, the cathode compartment was fed with 25 wt % Fe2O3 and was separated from the anode compartments by a GI-N417 reinforced membrane. The cell was externally heated to 83° C. at the anode compartments to maintain a temperature of ˜80° C. within the catholyte; two 5 cm2 Ni foils served as anodes evolving O2 and a 9 cm2 bifacial Cu plate served as the cathode. In between five consecutive electrolysis runs at 100 mA cm−2 (FIG. 26A), Fe was easily peeled from the cathode surface (FIG. 26B) and the collector was re-polished and re-inserted into the catholyte. The average ηFe was 91(3)% and was unaffected by repeated use of the cathode, whereas the steady-state cell voltage was observed to decrease from 2.45 to 2.31 V. Only a marginal amount of magnetic product was collected from the catholyte, which would have accounted for no more than 1% of the total ηFe if all the magnetic product was Fe, consistent with a reduction mechanism involving the electrocrystallization of Fe to a growing film.

Naturally occurring hematite was used as a reagent to assess the reaction scope for scalable sources of Fe. A representative, low-grade source of α-Fe2O3 containing 63 wt % Fe, 3 wt % Si, and 3 wt % Al was crushed by hand and sieved prior to electrolysis in the batch cell (FIGS. 27A-27C, Table 3). Neither the −150 mesh or −300 mesh α-Fe2O3 particles produced in this manner yielded selective Fe growth during 30 min electrolysis at J=100 mA cm−2 (ηFe=9% and 10%, respectively). In both cases, the morphology of the Fe product was small, ˜1 μm in diameter, in comparison to the solid particles fed to the cell which were 10 s of μm in diameter (FIGS. 28A-28B).

TABLE 3 Species Wt % Fe 62.6 Si 2.7 Al 2.5 C 30.6

The surface roughness of the Fe deposits differed from those obtained from purified α-Fe2O3. The lower ηFe observed for the 63 wt % Fe natural ore could be caused by the presence of impurities, hydrodynamic radius affecting particle adsorption and dissolution, or a combination of these and other factors. To investigate the role of naturally occurring metal oxide impurities, SiO2 and Al2O3 particles were blended with purified α-Fe2O3 as feedstocks for the chlor-iron cell. At J=100 mA cm−2, ηFe was inversely proportional to the quantity of SiO2 fed to the cell (FIG. 22C), although the polarization behavior was largely unaffected by SiO2. At an increased J of 300 mA cm−2, 8 wt % SiO2 in the catholyte led to an ηFe<37% during electrolysis (Table 4).

TABLE 4 SiO2 wt % Al2O3 wt % ηFe (%) ηFe vs. metal oxide concentration at 100 mA cm−2 0 0 94, 95 0 4 95, 95 8 0 58, 57 8 4 57, 54 ηFe vs. metal oxide concentration at 300 mA cm−2 0 0 87, 87 0 4 88, 93 8 0 36, 37 8 4 39, 41

The addition of Al2O3 at 4 wt % did not significantly impact ηFe for either J. Thermal oxidation of Fe(NO3)3 (>99% assay) at 500° C. in an open-air furnace followed by crushing and sieving produced larger and more continuous particles than the commercially available α-Fe2O3 particles (FIGS. 29A-29C), although excess iron oxide fines were not able to be separated from the larger domains. The −300 mesh fraction of these particles yielded ηFe=61% after 1 hour of electrolysis at J=100 mA cm−2, which was less than the ηFe obtained from purified Fe2O3 but greater than the ηFe obtained from crushed, low-grade ore.

Discussion

Fe and Cl2(g) were produced at J=100 mA cm−2 and η>90% from a single electrochemical cell, showing that the coupling of these anodic and cathodic reactions can be coupled to continuously generate Fe and Cl2 as separated products (FIGS. 16A, 16B, 17A, 17B, 18A, 18B, 25A-25E). Cells separated by a conventional cation-exchange membrane (FIG. 22A), a porous diaphragm (FIG. 23B), and a specialty bilayer cation exchange membrane (FIG. 25C) all yielded ηFe>90%, indicating that Cl− crossover through the cation-selective membrane was marginal or else had a marginal effect on the Fe-plating reaction. However, the chlorine evolution reaction was significantly affected by crossover through the separator affecting anolyte pH, and the use of a specialty bilayer membrane (N2030, Chemours) prevented crossover of the 7.5 mol/kg NaOH in the catholyte through the membrane to the dilute acid in the anolyte, required for collection of Cl2(g). The large pH gradient between the anolyte and catholyte led to an increase in the operating cell voltage (>3 V) relative to a cell producing only Fe and O2 from a single, alkaline electrolyte (1.7 V) (Leduc et al., J. Electrochem. Soc. 1959, 106:659; Allanore et al., J. Appl. Electrochem. 2010, 40:1957-1966).

A technoeconomic analysis of the levelized cost of electrolytic Fe—the cost of iron required to pay back capital costs, operating costs, and an assumed rate of return—was used to evaluate paths towards price-competitive green steel (Stinn et al., Electrochem. Soc. Interface 2020, 29:44; Manfield et al., “Economic Analysis of Air Pollution Regulations: Chlorine Industry, 2000). A similar capital cost to existing chlor-alkali plants was assumed, based on the similar operating conditions and bill of materials of the chlor-iron and chlor-alkali processes. The operating voltages and ηCl2 obtained in the 4 cm2 cell were used for the base case (FIG. 25B), and the sensitivity of the cost of iron to changes in η, the capital cost of the plant, the cost of iron ore, the cost of electric power, the cell voltage, and the cell lifetime are reported. The levelized cost of Fe is predicted to be $585/tonne Fe at existing industrial electricity rates ($75 MWh−1, U.S. grid average) but falls to $476/tonne if electricity can be procured at an average rate of $25 MWh−1. Because the cheapest and cleanest electricity is not available at all times of day, we explored the relationship between electricity cost, capacity factor, and levelized cost of iron. In the conceptual limit of free electricity, and all other variables maintained at the base case, Fe <$500/tonne is obtained for capacity factors >0.78, Fe at <$700/tonne is obtained only for plant capacity factors >0.56.

This work used a purified, aqueous NaCl feed to the anolyte, which could be sourced from abundant seawater after treatment to remove common impurities such as Mg2+ and Ca2+. The influence of these impurities on the anode process was not explored, as they are already removed from NaCl feeds to membrane chlor-alkali cells and the economics for the direct use of untreated seawater are unfavorable for industrial electrochemical processes (Small, ORE. JOM 1981, 33:67-75). Because the electrical energy required for the chlor-iron process (˜5 MWh/tonne Fe) is large compared to the energy required to purify seawater, even if the water were to be fully purified via reverse osmosis (0.003-0.004 MWh/tonne H2O), water is not expected to be a limiting resource in scaling this process.

The presence of SiO2 within the catholyte led to a substantially reduced ηFe (FIG. 22C), but did not appear to influence the polarization behavior of the cell (FIG. 22D) and elemental analysis did not detect Si within the film. Accumulation of dissolved silicates and aluminates in the catholyte should be avoided and could potentially adversely impact the long-term performance of the cathode or membrane. Whereas molten oxide electrolyzers operate with high selectivity in electrolytes containing SiO2, Al2O3, and MgO, it may be beneficial to first remove SiO2 from low-purity iron oxide ores to <2 wt % (e.g., with low-temperature electrolyzers) in order to yield catholyte conditions that enable ηFe>90% (FIG. 22C). Co-product NaOH may be used to pre-leach aluminates and silicate, which are soluble in concentrated alkali solutions, from the reactant feed.

High purity α-Fe2O3 prepared via calcination to form larger particles led to ηFe=61% during electrolysis (FIG. 29A) whereas commercial, high-purity α-Fe2O3 electrolyzed under similar conditions yielded ηFe>94%. The effective energy required for producing Fe from these two reactants is 5.3 MWh/tonne and 9.4 MWh/tonne, respectively. By comparison, typical grinding processes for ores require <0.1 MWh/tonne, suggesting that additional energy costs to reduce the size of the reactant particles will be warranted if they lead to modest increases to ηFe.

The Ti-supported RuOx anode used in this work enabled stable performance of the chlor-iron process device, with an average increase in voltage during a 4 hour electrolysis <3 μV hr−1, which is consistent with dimensionally stable anodes that last up to 10 years in service of chlor-alkali cells (Kempler et al., Energy & Env. Sci. 2020, 13:180 8-1817). Commercialized anodes intended for the chloro-alali process may yield similar efficiencies and improved durability, in the absence of significant crossover of OH− or dissolved metal oxides. The materials for anodic electrocatalyst films, current collectors, and membranes used in the disclosed exemplary chlor-iron electrochemical cells are commercially available. Moreover, the liquid electrolytes fed to the anode and cathode compartments are compatible with the chlor-alkali process, such that this method of ironmaking could be scaled along parallel manufacturing process or even as a retrofit of existing chlor-alkali process systems.

The generation of hydroxide ions at the cathode that are not consumed at the anode leads to a process that is a net-producer of a metal hydroxide as opposed to a net-consumer of metal hydroxides or acids, which must be leveraged against the final breakeven cost of the Fe. Separation costs associated with Fe and X2(g) are expected to be marginal. The typical purity of Cl2 gas leaving a chlor-alkali stack is 98-99% (remainder O2). In this example, although the generated NaOH includes metal/metal-oxide impurities, the low-purity NaOH could be used as a purifying reagent for commercial-grade iron ore (e.g., to leach out SiO2), or even used to capture and mineralize CO2, leading to a net-negative process for ironmaking thereby offsetting downstream emissions associated with steelmaking (Keith et al., Joule 2018, 2:1573-1594).

The market for primary Fe production is substantially greater than the existing market for Cl2. Assuming the chlor-iron process is capped by the present-day global market for Cl2(g) (˜100 MMT y−1) the limit to decarbonized Fe production via the chlor-iron process is ˜53 MMT y−1 which is 2-3% of the world's existing production. Nevertheless, due to the immense scale of the Fe market, the chlor-iron process could offset global CO2 emissions at a scale >100 MMT y−1. If electrochemistry is to play a major role in decarbonizing the zinc, copper, aluminum, lime, ammonia, ethylene, and iron industries, the majority of these emissions may be abated by cells evolving O2. Thus, the ultimate impact of the chlor-iron process may be maximized if it leads to a reduction in cell costs for direct iron oxide reduction and provides inexpensive streams of electrolyte to alkaline electrolyzers producing Fe and O2, which are not limited by offtake of the anodic product and could thus satisfy the remaining demand for primary Fe production.

The preferred morphology, purity, and throughput for Fe deposition may depend on the intended use of the Fe films. The gap between the cathode current collector and the bi-layer membrane establishes a practical limit for the thickness of iron films which can be produced from this method, where smaller gaps are preferred for reducing the area-specific resistance of the cell but larger gaps will reduce the risk of dendrites for extended growth periods. The morphology (particle diameter ˜10-50 μm) and purity (>95%) of the Fe produced is consistent with direct-reduced iron sponge used as a feedstock for electric arc furnaces (FIGS. 30A-30B, Table 5) (Raabe et al., Nature 2019, 575:64-74). Some aspects of the disclosed method could also be used to produce high surface area, monolithic iron electrodes for use in aqueous metal-air batteries.

TABLE 5 Species C O Fe wt % 0.3 1.4 98.3 atom % 1.4 4.7 93.9

In summary, embodiments of the disclosed electrochemical cell and process provided a scalable, efficient (˜5 MWh MTFe−1), and high-throughput (>100 mA cm−2) process for ironmaking that would produce zero or even net-negative direct greenhouse gas (GHG) emissions (FIG. 24). In some aspects, the overall cell reaction advantageously utilizes only low-cost iron oxide and seawater/brine as reactants. Fe was produced as easily collected free-standing films with >95 wt % purity, suitable for use in electric arc furnaces, such as furnaces that process scrap steel, and Cl2 was collected at industrially-relevant rates and selectivity. The co-production of NaOH is a distinguishing feature of this approach to ironmaking and could be used to purify low-grade iron ore to remove silicates or potentially used in a downstream process for CO2 capture and mineralization, leading to a net-negative GHG emission ironmaking process with the potential to directly abate up to 120 MMT y−1 CO2 emissions from blast furnaces.

Example 4

Hydrogen co-production increases when the metal oxide particles are larger and/or less porous. In one trial, precipitated Fe2O3 particles (SG_Fe2O3, FIG. 31A) having a surface area of 65 m2g−1 were fed to the cell and yielded 76% faradaic efficiency at −1.2 V (average current density −95 mA cm−2) towards Fe with the remaining current assumed to go towards the hydrogen evolution reaction (HER). In another trial, annealed Fe2O3 particles (SO_Fe2O3, FIG. 31B) having a surface area of 4.5 m2g−1 were fed to the cell and yielded 8% faradaic efficiency towards Fe at −1.2 V (average current density −54 mA cm−2) with the remaining current assumed to go towards the HER. The obtained result was compared to particles from FIGS. 16A and 16B (TF_Fe2O3) that have 13 m2g−1 surface area. The results are shown in FIG. 31C.

Hydrogen co-production also increases as the current density increases. Each of the particle types was fed to the cell at increasing potential and collected data was presented as relation of passed voltage to partial current density towards Fe production jFe (FIG. 31C). This led to the increased current density and reduced faradaic efficiency towards iron production with the remaining current assumed to go towards the HER.

In view of the many possible embodiments to which the principles of the disclosed invention may be applied, it should be recognized that the illustrated embodiments are only preferred examples of the invention and should not be taken as limiting the scope of the invention. Rather, the scope of the invention is defined by the following claims. We therefore claim as our invention all that comes within the scope and spirit of these claims.

Claims

1. An electrowinning process, comprising:

providing an electrochemical cell comprising (i) a cathode comprising low-carbon steel, copper, iron, graphite, vitreous carbon, or titanium, (ii) an anode comprising an oxide coating comprising Ru, Pt, Ir, or any combination thereof, on a conducting substrate, (iii) a separator between the cathode and the anode, the separator comprising a porous composite or a cation-selective membrane, and (iv) a voltage source electrically connected to the cathode and the anode;
providing a catholyte comprising (i) water, (ii) a metal hydroxide comprising Q, where Q is an alkali metal, an alkaline earth metal, or a combination thereof, and (iii) suspended metal ore particles comprising MxOy where M is a metal and x and y are integers;
providing an anolyte comprising water and a halide salt comprising Q and X, where X is Cl, Br, or a combination thereof; and
applying a voltage across the electrochemical cell to effect (i) reduction of the MxOy in the catholyte to provide the metal M and additional metal hydroxide comprising Q, and (ii) production of O2, X2, XO−, or any combination thereof in the anolyte.

2. The electrowinning process of claim 1, wherein:

the anolyte has a bulk average pH≤3 and X2 production is greater than O2 production; or
the anolyte has a bulk average 3<pH<7 and XO− production is greater than O2 and/or X2 production; or
the anolyte has a bulk average pH≥7 and O2 production is greater than X2 production.

3. The electrowinning process of claim 2, wherein:

the anolyte has a pH≤3 and the anode is in direct contact with the separator; or
the anolyte has a pH≥7 and the anode is spaced apart from the separator.

4. The electrowinning process of claim 1, applying the voltage across the electrochemical cell further effects production of H2 in the catholyte.

5. The electrowinning process of claim 1, wherein:

(i) M is Fe, Mn, Ni, Cr, Co, Zn, or any combination thereof; or
(ii) Q is Li, Na, K, Rb, Cs, Mg, Ca, or any combination thereof; or
(iii) both (i) and (ii).

6. The electrowinning process of claim 1, wherein:

(i) the metal ore particles comprise Fe2O3; or
(ii) Q is Na;
(iii) X is Cl; or
(iv) any combination of two or more of (i), (ii), and (iii).

7. The electrowinning process of claim 1, wherein:

(i) the catholyte comprises from 50 g/L to 500 g/L of the suspended metal ore particles prior to applying the voltage; or
(ii) the catholyte comprises from 10 wt % to 50 wt % of the metal hydroxide prior to applying the voltage; or
(iii) the anolyte comprises from 10 wt % to 50 wt % of the halide salt prior to applying the voltage; or
(iv) any combination of two or more of (i), (ii), and (iii).

8. The electrowinning process of claim 1, further comprising:

(i) continuously or periodically removing X2 and/or O2 generated in the anolyte; or
(ii) periodically removing at least a portion of the metal M from the cathode; or
(iii) continuously or periodically removing H2 generated in the catholyte; or
(iv) any combination of two or more of (i), (ii), and (iii).

9. The electrowinning process of claim 1, further comprising:

(i) periodically adding a quantity of the metal ore particles to the catholyte; or
(ii) periodically adding a quantity of the halide salt to the anolyte; or
(iii) both (i) and (ii).

10. The electrowinning process of claim 1, wherein the metal ore particles are obtained from a metal ore feedstock further comprising aluminates, silicates, or aluminates and silicates, the method further comprising leaching at least a portion of the aluminates, silicates, or aluminates and silicates from the metal ore feedstock by contacting the metal ore feedstock with a hydroxide solution to provide the metal ore particles.

11. The electrowinning process of claim 10, wherein the hydroxide solution is a spent catholyte obtained from the electrochemical cell after applying the voltage across the electrochemical cell.

12. The electrowinning process of claim 1, wherein the anolyte comprises concentrated seawater having a halide salt concentration of from 10 wt % to 50 wt %.

13. The electrowinning process of claim 1, wherein:

providing the electrochemical cell further comprises providing a cell stack comprising (i) a number of the electrochemical cells, a cathode electrical connector connecting cathodes of each of the electrochemical cells in parallel, an anode electrical connector connecting anodes of each of the electrochemical cells in parallel, and a voltage source electrically connected to the cathode electrical connector and the anode electrical connector, or (ii) a number n of the electrochemical cells, a number n−1 of conductive bipolar plates wherein a conductive bipolar plate is positioned between each adjacent pair of electrochemical cells, a cathode electrical connector connected to a cathode of a first electrochemical cell in the series, an anode electrical connector connected to an anode of a last electrochemical cell in the series, and a voltage source electrically connected to the cathode electrical connector and the anode electrical connector;
providing the catholyte further comprises providing the catholyte within each electrochemical cell of the cell stack;
providing the anolyte within the anode compartment further comprises providing the anolyte within each electrochemical cell of the cell stack; and
applying a voltage across the electrochemical cell further comprises applying the voltage across the cell stack to effect reduction of the MxOy in each cathode compartment to provide the metal M and formation of Cl2 gas in each anode compartment.

14. The electrowinning process of claim 1, wherein:

(i) the voltage applied is from 2 V to 5 V per electrochemical cell; or
(ii) the electrochemical cell or cell stack is operated at a current density of from 20 mA cm−2 to 500 mA cm−2; or
(iii) the electrochemical cell or cell stack is operated at a temperature from 25° C. to 150° C.; or
(iv) any combination of (i), (ii), and (iii).

15. An electrochemical cell, comprising:

a cathode comprising low-carbon steel, copper, iron, graphite, vitreous carbon, or titanium;
an anode comprising an oxide coating comprising Ru, Pt, Ir, or any combination thereof, on a conducting substrate; and
a separator between the cathode and the anode, the separator comprising a cation-selective membrane that is permeable to alkali metal cations, alkaline earth metal cations, or a combination thereof.

16. The electrochemical cell of claim 15, further comprising:

a second anode; and
a second separator between the cathode and the second anode, the second separator comprising a cation-selective membrane that is permeable to alkali metal cations, alkaline earth metal cations, or a combination thereof.

17. The electrochemical cell of claim 15, further comprising:

(i) gas collecting means for collecting gas generated at the anode; or
(ii) gas collecting means for collecting gas generated at the cathode; or
(iii) a magnet operable to be passed over a surface of the cathode; or
(iv) a catholyte mixing means and an anolyte mixing means; or
(v) a voltage source electrically connected to the cathode and the anode; or
(vi) any combination of two or more of (i), (ii), (iii), (iv), and (v).

18. The electrochemical cell of claim 15, further comprising:

a catholyte comprising (i) water, (ii) a metal hydroxide comprising Q, where Q is an alkali metal, an alkaline earth metal, or a combination thereof, and (iii) suspended metal ore particles comprising MxOy where M is a metal and x and y are integers; and
an anolyte comprising water and a halide salt comprising Q and X where X is Cl or Br.

19. A cell stack, comprising:

(a) a number n of electrochemical cells according to claim 15, a cathode electrical connector connecting cathodes of each of the electrochemical cells in parallel, and an anode electrical connector connecting anodes of each of the electrochemical cells in parallel; or
(b) a number n of electrochemical cells according to claim 13 arranged in series, a number n−1 of conductive bipolar plates, a conductive bipolar plate positioned between each adjacent pair of electrochemical cells, a cathode electrical connector connected to a cathode of a first electrochemical cell in the series, and an anode electrical connector connected to an anode of a last electrochemical cell in the series.

20. The cell stack of claim 19, further comprising:

a catholyte within each electrochemical cell, the catholyte comprising (i) water, (ii) a metal hydroxide comprising Q, where Q is an alkali metal, an alkaline earth metal, or a combination thereof, and (iii) suspended metal ore particles comprising MxOy where M is a metal and x and y are integers; and
an anolyte within each electrochemical cell, the anolyte comprising water and a halide salt comprising Q and X where X is Cl or Br.
Patent History
Publication number: 20240200217
Type: Application
Filed: Feb 20, 2024
Publication Date: Jun 20, 2024
Applicant: University of Oregon (Eugene, OR)
Inventors: Paul Kempler (Eugene, OR), Anastasiia Konovalova (Eugene, OR), Shannon W. Boettcher (Eugene, OR), Berkley B. Noble (Alameda, CA)
Application Number: 18/582,228
Classifications
International Classification: C25C 1/06 (20060101); C25B 1/16 (20060101); C25C 1/02 (20060101);