DEVICE AND METHOD FOR CONVERSION OF CARBON DIOXIDE TO HYDROCARBON MOLECULES

Description

CROSS-REFERENCE TO RELATED APPLICATIONS

This application claims the benefit of priority from U.S. Provisional Application Ser. No. 63/700,880 filed on Sep. 30, 2024 and entitled “Device and Method for Conversion of Carbon Dioxide to Hydrocarbon Molecules,” the entire contents of which are incorporated herein by reference for all purposes.

DEPARTMENT OF ENERGY FUNDING

This invention was made with Government support under contract number DE-SC0023939 awarded by DOE, Office of Science. The Government has certain rights in this invention.

FIELD OF THE INVENTION

The embodiments of the present invention are directed to electrochemical systems and their methods of use. In particular, various embodiments are disclosed directed to electrochemical systems comprising thin, porous metal sheets for conversion of carbon dioxide into hydrocarbon chemicals and fuels.

BACKGROUND OF THE INVENTION

A sudden increase of carbon dioxide (CO₂) in the atmosphere is primarily caused by the burning of fossil fuels (coal, oil, and natural gas) for energy and industry, which releases carbon stored for millions of years into the atmosphere at an unprecedented rate. One solution to this problem is to minimize and eliminate usage of fossil-based fuels and products. Many CO₂ capture technologies have been developed to produce pure CO₂ from carbon emission sources and from atmosphere directly. In particular, direct air capture (DAC) of CO₂ may supply CO₂ wherever CO₂ may be disposed of or utilized not limited by transportation and storage.

Economic CO₂ conversion technologies are desired to produce hydrocarbon fuels and/or chemical products using renewable electricity. For recent years, production cost of the renewable electricity, such as solar photovoltaic (PV) and wind turbine generator, has come down rapidly. Storage and transportation have become a bottleneck with widespread usage of renewable electricity. Electrochemical conversion of CO₂ to hydrocarbon products may solve both CO₂ conversion and electricity storage problems.

Many electrochemical CO₂ conversion technologies have been reported, which may be classified into indirect conversion and direct conversion routes of CO₂ and H₂O. In the indirect conversion route, hydrogen gas is produced first by water electrolysis using renewable electricity and then, used to react with CO₂ and/or CO to form hydrocarbon products, such as pure hydrocarbons and oxygenates (methanol, ethanol, etc.). The indirect conversion also includes reduction of CO₂ into CO. In the direct conversion route, CO₂ and H₂O are directly converted into pure hydrocarbons (methane, ethylene, etc.) and oxygenates (formic acid, methanol, carboxylic acids, glycols, ethanol, etc.). The direct conversion route avoids usage of water electrolyzer and is desirable to have a simple, efficient, and low-cot CO₂ conversion unit. A lot of scientific studies and inventive ideas have been reported in research and patent publications in recent years. However, challenges remain to make the CO₂ conversion process become economically competitive with current products. For example, electrical efficiency is too low and electrochemical (EC) cell costs are too high.

SUMMARY OF THE INVENTION

Various embodiments disclosed herein may provide a low-temperature (T) and low-pressure (P) unit to reduce the equipment cost, transportation and installation cost, and operation cost. The low-T and low-P units may be manufactured at low cost as compact and modular units. The modular unit may be readily transported and installed at the production site. The low-T and low-P modular unit renders quick startup up and shutdown to reduce the unit turn-around costs. It is known from current chemical processing industries that it takes a lot of time to heat up and cool down a large high-temperature reaction unit and takes a lot of time to pressurize and depressurize a large high-pressure reaction unit. The quick unit turn-around is needed to efficiently utilize renewable electricity that may fluctuate with time and climate conditions.

BRIEF DESCRIPTION OF THE DRAWINGS

FIG. 1A illustrates a single cell electrochemical device for conversion of CO₂; FIG. 1B illustrates a multicell electrochemical device for conversion of CO₂.

FIG. 2A illustrates an open view of basic components of the cell stacking in a gas/liquid two-phase electrochemical reactor for conversion of CO₂; FIG. 2B illustrates an open view of basic components of the cell stacking in a gas/liquid two-phase electrochemical reactor for conversion of CO₂ further including a current and/or pressure distributor.

FIG. 3A is a configuration of anodic catalytic reaction zone including a Porous Nickel alloy sheet; FIG. 3B is a configuration of anodic catalytic reaction zone including an anodic reaction zone of graded pore structures; FIG. 3C is a configuration of anodic catalytic reaction zone including an anodic catalyst coating on gas diffusion layer.

FIG. 4A is a configuration of membrane separator having a porous ceramic coating on porous Nickel alloy sheet; FIG. 4B is a configuration of membrane separator having a polymer membrane/ceramic coating/porous Nickel alloy sheet; FIG. 4C is a configuration of membrane separator having a polymer membrane on porous Nickel alloy sheet.

FIG. 5A is a configuration of cathodic catalytic reaction zone having a catalyst coating on compressible gas diffusion layer; FIG. 5B is a configuration of cathodic catalytic reaction zone having a catalyst coating on the ceramic membrane surface and encapsulated by gas diffusion layer; FIG. 5C is a configuration of cathodic catalytic reaction zone having a catalyst coated inside pores of a porous cathode support sheet.

FIG. 6A is a carbon catalyst support containing micro and mesopores; FIG. 6B is a carbon catalyst support containing micro and mesopores having immobilization of nano catalyst particles and electrolyte solutions.

FIG. 7A illustrates a structure of zirconia coated on porous Nickel sheet showing a fractured wall; FIG. 7B illustrates a structure of zirconia coated on porous Nickel sheet showing a surface of bare porous Nickel alloy sheet; FIG. 7C illustrates a structure of zirconia coated on porous Nickel sheet showing a surface of zirconia membrane.

FIG. 8 illustrates the permeation rate of humid air through a zirconia/Nickel sheet at different pressure gradient at room temperature.

FIG. 9A illustrates the water electrolysis activity of zirconia/nickel sheet at room temperature of Nickel (+)/ceramic-coated Nickel (−); FIG. 9B illustrates the water electrolysis activity of zirconia/nickel sheet at room temperature of Nickel (−)/ceramic-coated Nickel (+).

FIG. 10 is a process flow diagram of laboratory test setup for conversion of CO₂ and H₂O into C₂₊ oxygenated fuels.

FIG. 11A illustrates a cell assembly for experimental testing having a cover plate with built-in flow channels; FIG. 11B illustrates a test cell assembly comprising different layers.

FIG. 12 illustrates a product distribution of a liquid sample (#7) from condensation of cathode effluent gas.

FIG. 13 illustrates ethanol detection in cell cathode effluent.

FIG. 14 illustrates the impact of CO₂ gas purity on electrochemical cell productivity.

FIG. 15 illustrates the co-feed of liquid and gas to cathode side.

FIG. 16 illustrates the flow through cell.

FIG. 17A illustrates a layered view of the cell stack of an electrochemical cell of larger working area; FIG. 17B illustrates the inlet and outlet ports on the anode side end plate of the cell stack of an electrochemical cell of larger working area.

FIG. 18 illustrates the impact of feed CO₂ gas humidification on gaseous product formation.

FIG. 19A illustrates a layered view of a two-cell stacking; FIG. 19B illustrates the inlet and outlet ports on the anode side end plate.

FIG. 20A illustrates the components of a unit cell plate; FIG. 20B illustrates a view of the cathode side of the unit cell; FIG. 20C illustrates a view of the anode side of the unit cell.

FIG. 21A illustrates the design features of cathode flow spacer; FIG. 21B illustrates the design features of anode flow spacer.

FIG. 22 illustrates a stack of unit cells into a reactor module.

FIG. 23 illustrates a process flow diagram of the electrochemical (EC) conversion plant of carbon dioxide and water into hydrocarbons.

FIG. 24 illustrates the integration of direct air capture and electrochemical conversion into one system.

FIG. 25 illustrates the production of carbon-neutral hydrocarbon products using solar electricity.

FIG. 26A illustrates Modular DACC units for easy transport and installation of a modular AHX cart, air tunnel and regeneration chambers; FIG. 26B illustrates Modular DACC units for easy transport and installation of a modular EC reactor and product separation units.

DETAILED DESCRIPTION OF THE INVENTION

Various embodiments disclosed herein provide a new type of EC reactor to address the above mentioned problems with related EC reactors. Furthermore, the EC reactor used in various embodiments disclosed herein may allow for manufacturing of low-cost, compact module units that can be readily transported and deployed, in particular to operating sites where strand electricity is available. The EC reactor of this invention also facilitates integration of CO₂ conversion with DAC units by reducing energy consumption and capital costs of both CO₂ capture and conversion processes.

Various features of the EC reactor device used in various embodiments disclosed herein may conduct the above reaction are illustrated with a single cell in FIG. 1A. The device comprises 1) one end plate to allow liquid to be introduced into liquid flow channels and discharge the liquid and anodic product out of the liquid flow channel, 2) a liquid flow channel to allow uniform distribution of the liquid feed over anodic reaction zone, 3) a anodic reaction zone that is positively charged to conduct anodic reaction, 4) a membrane separator that allows hydroxide ions to permeate while blocking crossover of cathodic gas and anodic liquid, 5) a cathode reaction zone that is negatively charged to conduct the cathodic reaction, 6) a gas flow channel that distributes the gas feed over the cathodic reaction zone uniformly and removes the cathodic reaction product, and 7) an end plate to allow introduction of feed gas into the gas flow channel and discharge of the cathodic product out of the gas channel.

Various embodiments disclosed herein may comprise two or more cells as illustrated in FIG. 1B. One set of the cathodic reaction zone/membrane separator/anodic reaction zone is defined as one full cell. In the multicell device, only the liquid flow channel and gas flow channel next to the end plate serves one full cell. The liquid or gas flow channel in rest of the cell stack may be shared with two adjacent cells. As shown in FIG. 1B, the liquid flow channel (2) may be disposed of between two identical anodic reaction zones (3), and the gas flow channel (6) may be disposed of between two identical cathodic reaction zones (5). 3 cells are shown in FIG. 1B. With the standardized components and stacking method, more numbers of the cell may be stacked in one device. In the multicell stack, gas flow conduits are built into these components to allow feed gas flow from the cathode gas inlet to be distributed into all the gas flow channels in the device and allow collection of all the cathodic reaction products out of all the gas channels into the cathode gas outlet. Similarly, the liquid flow conduits built into these components allow the liquid feed flow introduced from the anodic liquid inlet to be distributed into all the liquid flow channels in the stack uniformly and collection of the cathodic reaction product from all the gas flow channels to the cathode outlet. All the cells are applied with same voltage difference between the positive and negative reaction zones.

CO₂ and H₂O vapor may be introduced into cathodic reaction zone to form hydrocarbon products (Cp) and hydroxide ions through the following electrochemical reaction:

The hydroxide ions move across a membrane separator into anodic reaction zone to be reduced into oxygen gas and water through the following electrochemical reaction:

In various embodiments disclosed herein, the anode may be positively charged to take electrons away from the electrochemical reaction, while the cathode may be negatively charged to provide electrons for the electrochemical reaction.

The liquid feed in the anode may be preferably to be a water-based solution of pH greater than 7, i.e., alkaline electrolyte. The gas feed in the cathode contains CO₂ and water vapor with relative humidity greater than 0 and less than 100%. The device may be operated with pressure in the gas channel, always greater than pressure in the liquid channels to avoid crossover of gas and/liquid. The device may be operated at low pressures with the liquid channel pressure being in a range of 1 to 2 bars (absolute) and gas channel pressure being in a range of 1.01 to 3.0 bars (absolute). Most EC cells reported in the literature are operated at room temperature. By contrast, various embodiments disclosed herein may be operated between room temperatures and water boiling point, in the range of 20 to 100° C. The device is preferably operated at temperatures above 35° C. so that no additional refrigeration cooling is needed in warm climates and/or waste heat released from the device may be recovered for heating values. Increasing operation temperature above room or environmental temperature may enhance productivity and energy efficiency, and reduce the EC reactor and product separation costs. At elevated reaction temperatures, more H₂O reactant may be fed into the cathode reaction channel along with CO₂ gas, because related reactor designs do not allow presence of liquid water in the cathode gas channels. Product separation cost and energy consumption from cathode and/or anode reactor effluents may be reduced by increasing temperature. However, water boils at 100° C. under atmospheric pressure. To keep water-based electrolyte liquid stay in liquid phase at reaction temperatures above 100° C., the reactor needs to be pressurized, which will increase capital costs. Thus, the upper operating temperature is preferably below 120° C. to operate the reactor under nearly atmospheric or low pressures. The voltage applied to each cell between the cathode and anodic reaction zones may be in the range of about 1.0 to about 3.0V.

The design features of the EC reactor or device of this invention are further illustrated with open-up views in FIGS. 2A and 2B. FIG. 2A shows one kind of cell configuration with direct contacting of the flow channels and electrode reaction zone. The cathodic flow channel (6) may be made of electron-conducting material with hydraulic diameter in the range of 0.3 to 5.0 mm. The flow channels are defined by solid channel walls and the cathode reaction zone. The channel wall may be negatively charged to provide electrons to the cathode reaction zone (3). CO₂ and H₂O molecules may diffuse from the gas channel into the reaction zone to react to electrons over a cathode catalyst to form desired reaction products. The cathode reaction product may diffuse from the reaction zone into the cathode flow channel. The cathodic reaction zone that is measured from gas/cathode interface to the cathode/membrane interface may be preferably of a depth (or thickness) in a range from 0.2 to 4.0 mm and of porosity from 0.2 to 0.90. Similarly, the anodic flow channel (2) may be made of electron-conducting material with hydraulic diameter in the range of 0.3 to 5.0 mm. The flow channels are defined by solid channel walls and the anode reaction zone (3). The channel wall may be positively charged to remove electrons released from the anodic reaction. Anodic reaction product—O₂ gas diffuses from the anodic reaction zone into the anode flow channel. Heat generated from the reaction process may be conducted to the anode flow channel. The anodic reaction zone that is measured from liquid/anode interface to the anode/membrane interface is preferably of depth (or thickness) that ranges from 0.0.040 to 2.0 mm and of porosity from 0.2 to 0.80.

FIG. 2B shows another kind of cell configuration with an electron and/or pressure distributor being placed between the flow channel and the reaction zone. The distributor may distribute electrons uniformly onto whole area of the reaction zone and also distributes compression force uniformly onto the reaction zone so that all the components in the cell stack may be compressed to have intimate contacts by the two end plates. Because electrons and hydroxide ions cannot be conducted in a void space, the cathode reaction zone/membrane/anodic reaction zone may preferably be in intimate contact to ensure continuous conducting paths for electrons and hydroxide ions. Without the distributor, compression force applied from the end plates may not be transferred to the flow channel/reaction zone interfaces to compress these components into intimate contact. The distributor may have high mechanical rigidity to transfer the compression force from the channel wall to the channel area, high electron conductivity, and high permeance without imposing transport resistance of reactants and/or products between the reaction zone and flow channel. The distributor materials include 1) metal meshes such as stainless steel, 2) perforated metal plates, and 3) carbon fiber composites. Thickness and porosity of the distributor is preferably to be 0.02 to 0.5 mm and 0.30 to 0.70, respectively.

The anodic reaction zone is essential to conduct anodic reactions efficiently. Its structure features are illustrated with three configurations in FIG. 3A-3C. In FIG. 3A, the anodic reaction zone comprises a porous Nickel alloy sheet with thickness ranged from 0.025 to 0.20 mm, porosity of 0.35 to 0.65, and pore sizes from50 nm to 5000 nm. This kind of the porous Nickel alloy sheet has been invented and developed by Molecule Works Inc. Their structures and preparation methods were disclosed in prior art. The porous Nickel alloy by itself may be an active electrode to catalyze the anodic reaction. The porous Nickel sheet may be further modified to increase its anodic reactivity for oxygen evolution. The porous Nickel sheet provides smooth surface textures to support a thin membrane separator.

As shown in FIG. 3B, the above porous Nickel sheet may be added with another layer of electron and thermal conductive material of larger pores, larger porosity, and/or more compressibility. The additional layer has a thickness of 0.05 to 1.0 mm, pore sizes from 10 to 200 μm, and porosity from 0.5 to 0.98. This additional layer provides more compressibility than the porous Nickel sheet and facilitates removal of oxygen gas. The candidate materials may be foam and meshes of Nickel alloy and stainless steel.

The anodic reaction zone in the present EC cell serves a function like anodic electrode in alkaline water electrolyzer. Thus, anodic electrode material used in the water electrolyzer may be applied here. FIG. 3C shows an example of coating of anodic catalysts such as NiFe onto a compressible gas diffusion layer of thickness ranged from 0.02 to 2.0 mm. The compressible gas diffusion layer may be foam and meshes of Nickel alloy and stainless steel.

In addition to high conductivity for hydroxide ions and high electrical insulation properties, the membrane separator for the present EC is stable in the hydrocarbon environment and at elevated temperatures up to 100° C. The hydrocarbon products, such as methanol, ethanol, acetone, and propanol, tend to react with polymeric membranes. Polymeric material often has a low softening temperature and is chemically and thermally less durable than the ceramic membrane. FIGS. 4A-4C show examples of the membrane separator configuration for the present EC device.

FIG. 4A shows a porous ceramic membrane coated on the porous metal sheet. This membrane has been invented and developed by Molecule Works Inc. The membrane may be made of durable ceramic-type materials such as zirconia. The membrane may have a thickness of 5 to 40 μm above the porous metal surface. The membrane may have pores much smaller than the pores on the metal support sheet, preferably in the range of 2 to 100 nm. The ceramic membrane layer may provide electrical insulation and immobilization of the electrolyte solution by capillary force as described by the following Young-Laplace equation:

Δ ⁢ p c = 2 ⁢ γcos ⁡ ( θ ) r c ( 4 )

where Δp_c=capillary pressure; γ is the interfacial tension of gas/liquid; θ=contact angle of the liquid on the solid; r_c=radius of pore curvature.

The contact angle is preferably less than 90° or about zero. The capillary pressure increases with decreasing radius of pore curvature. For aqueous electrolyte solutions and zirconia membrane, the contact angle is about zero and capillary pressure over a temperature range of 25 to 100° C. is estimated to be about 1.2 to 1.4 bar for the pore radius of 1 μm and 12 to 14 bar for the pore radius of 0.1 μm.

According to the Hagen-Poiseuille's law, under a pressure gradient, the electrolyte fluid permeates through the porous membrane from anode to cathode side:

J F = P m ⁢ Δ ⁢ p m δ m ( 5 )

Where J_F=fluid permeation flux, m³/(m²·s); P_m=membrane permeability, m²/(Pa·s); Δp_m=pressure gradient of fluids across the membrane, Pa; δ_m=membrane thickness, m.

In the present reactor, the gas flow channel pressure (p₁) is maintained higher than the liquid flow channel pressure (p₂), negative pressure gradient stops permeation flow of the electrolyte liquid from the anode toward the cathode side.

When p₁>p₂, there is a positive pressure gradient for gas to permeate from the cathode toward the anode side. However, gas permeation from the gas flow channel into the liquid flow channel is blocked by the electrolyte immobilized inside the ceramic membrane pore.

When the gas-to-liquid pressure gradient, p₁-p₂, is maintained less than the capillary pressure, Δp_c, the liquid inside the membrane pore would not be pushed out.

Given the cell designs of geometries and materials, p₁ and p₂ are the process control conditions during cell operation. Pressure in the anodic side is preferred to be 1 (atmospheric pressure) to 3 bar, preferably 1 to 1.5 bar. The cathode side is maintained at least 1 kPa higher than the anode side, preferably 5 kPa greater.

As described by the above equation, Δp_c is determined by the membrane pore size and distribution, membrane material, and electrolyte solution. The electrolyte and membrane material are chosen to build the cell in such a way that the pore is wettable over the possible ranges of process operation conditions (T, compositions). The preferred membrane material is zirconia, alumina, ceria, silicon carbide, and their mixture. Other suitable membrane materials are within the contemplated scope of disclosure. The ceramic membrane may be doped or impregnated with certain additives to promote its sintering at lower temperature and/or its wettability to the electrolyte solution. The membrane pore is preferably with mean pore size less than 0.2 μm, substantially free of pores greater than 1 μm.

Membrane thickness may be another design factor that impacts operation and effectiveness. To minimize the ionic transport resistance, the membrane is preferred to be as thin as possible. To eliminate electrical short-circuits, the membrane is desired to have a certain thickness for uniform coverage of all the metallic support surface. To minimize formation of membrane defects, such as delamination and cracks, the membrane is preferred to be as thin as possible. Based on experimental tests, the membrane thickness is preferred to be 5 to 40 μm, more preferably 5 to 25 μm.

The metallic support material may be nickel and its alloys, titanium, steel, and stainless steel. The support thickness is preferred to be less than 0.2 mm to reduce the material cost and weight. For alkaline-based electrochemical cells, nickel and its alloy are preferred. The nickel and its alloy are known as electrode materials for water electrolysis. The anodic reaction of the present electrochemical cell is the same as the anode of alkaline water electrolysis cells. With the Nickel-type support material, the membrane and anodic electrode are integrated as one structural body as membrane electrode assembly (MEA). The porous metal support sheet has smaller pore sizes than other metal meshes, metal foams, and sintered metals, which are used in current electrochemical cells or devices. The porous metal support sheet is preferred to have mean pore size less than 2 μm, substantially free of pores greater than 5 μm. The porosity is preferred to be 0.35 to 0.55.

To safe guide electrical insulation function of the thin ceramic membrane, a thin polymer membrane may be laid on the ceramic membrane surface to cover possible defects that may result in short circuiting or electron conduction, as illustrated in FIG. 4b. In this configuration of polymer/ceramic composite or hybrid membrane, the polymeric membrane primarily provides additional electrical insulation. The polymer membrane should be as thin as possible. Its thickness is preferably in the range of about zero to about 0.2 mm. The polymer membrane can be an alkaline ion exchange membrane or porous membrane with high conductivity for hydroxide ions.

When a chemically and thermally stable polymeric membrane is available, the polymeric membrane may be directly laid out on a porous Nickel alloy sheet as shown in FIG. 4c. The porous Nickel sheet provides mechanical strength and surface smoothness to support the thin polymer membrane. The porous Nickel sheet may also function as the anodic electrode.

The cathodic reaction zone has a large impact on productivity, selectivity, and electrical efficiency of CO₂ and H₂O conversion processes. The essential structure features in this zone comprises 1) vapor-phase (or gas-phase) pores for rapid diffusion of CO₂ and H₂O from the gas flow channel to the catalytic site, 2) the catalytic sites for reaction of CO₂ and H₂O with electrons to form desired hydrocarbon product molecules and hydroxide ions, 3) electron conducting paths to transfer electrons from the distributor to the active sites, and 4) conducting paths for hydrated hydroxide ions to transport to and through the membrane separator from the catalytic site. Fabrication of the cathodic reaction zone is illustrated with three types of possible structure in FIGS. 5A-5C.

As shown in FIG. 5A, a porous cathode catalyst is coated on a porous carbon-support sheet. The catalyst coating thickness is less than about 50 μm. The thicker coating would cause cracks or delamination. The catalyst coating is uniform enough to obtain intimate and uniform contact with the membrane surface in the device. For example, the surface roughness of the catalyst coating should be less than 10 μm. The coating can be made by dry coating processes such as physical vapor deposition and chemical vapor deposition or by wet chemistry coating processes, such as spray coating and casting. The porous carbon sheet has high porosity and large pores for rapid diffusional mass transfer of reactants and products. The porosity and pore size are in the range of 0.30 to 0.90 and 10 to 500 μm, respectively. The porous carbon sheet has high electron and thermal conductivity. Its areal electrical resistance is preferably less than 1.0 Ωcm², that is, the voltage loss is 1.0×0.1=0.1 V at current density of 100 mA/cm². Its thermal conductivity is preferably greater than 1 W/(K·M). The porous carbon sheet needs to be thick enough to have mechanical strength and controlled degree of compression over a range of 5 to 50% during the cell assembly. The sheet is thin enough to minimize mass and heat transfer resistance. The thickness is preferably 0.1 to 3.0 mm. Examples of such porous carbon sheets are carbon paper and carbon felt, which are used in fuel cells as a gas diffusion layer.

FIG. 5B shows a cathode reaction zone that may accommodate a thicker catalyst layer. In this configuration, the cathode catalyst layer is sandwiched between the membrane separator and the porous carbon sheet with chemical and physical properties as described above. In this way, the catalytic layer thickness may be controlled over a range of about 0.010 to about 1.0 mm. The catalyst layer may be made by wet chemistry methods, such as printing, casting, painting, and laminating. In this configuration, the cathode reaction zone/membrane/anodic electrode may be made as one integrated sheet or plate to lower manufacturing costs and render the cell assembly.

To increase the reactant/catalyst/electron multiphase contacting area per unit volume of the cathode reaction zone, nano cathode catalyst may be deposited inside pores of the porous carbon sheet. As illustrated in FIG. 5C, the nano cathode catalyst may be coated on the carbon fiber surface inside a porous carbon sheet with the properties as described above. Volume fraction of the catalyst inside the sheet is small relative to the sheet porosity, preferably in the range of about 0.01 to 0.2, to have insignificant impacts on mass transfer rates of the reactants and products. The catalyst may be loaded in a fraction of the porous carbon sheet to the whole sheet thickness.

The cathode catalyst to produce C₂₊ hydrocarbons and oxygenates is preferably copper with primary particle sizes from 2 to 200 nm, which may be characterized by surface areas of 10 to 200 m²/g. The metallic copper may be promoted with some other elements, such as carbon, Sn, Zn and copper oxide. Binders may be incorporated into the copper catalyst coating or layer to modify surface hydrophobicity and enhance immobilization of the copper catalyst. One example of such a binder is polytetrafluoroethylene (PTFE), which has excellent chemical and thermal stability under the present EC cell reaction conditions. Alkaline conducting materials may also be added to the cathode catalyst layer to enhance conductivity of the hydrated hydroxide ions. Examples of such a conducting material are alkaline-conducting polymers, hydroxide ion-exchanging resins, carbonates, bicarbonates, hydroxides, and their mixture. The catalyst coating or layer may contain electron-conducting materials to assure continuous electron transfer paths throughout the whole catalyst layer. Examples of electron conducting materials are carbon, graphite, graphene, carbon nanotubes, or their mixtures. The active catalyst in the catalyst coating or layer may have a volume fraction of about 0.1 to 0.7.

In porous carbon (including graphite) papers, carbon felts, or carbon sheets, the carbon typically exists in a dense fiber form with a diameter ranged from 2 to 100 μm, mostly within 5 to 10 μm. The carbon particles with sizes in a range of 2 to 100 μm also exists as a dense matter. To enhance immobilization of the catalyst particles and/or electrolyte solution, micro or mesopores may be created on exterior surface of the solid carbon support in either particle or fiber forms. As illustrated in FIG. 6A, the micro and mesopore sizes are substantially less than the carbon support size. The pore opening is preferably in a range of approximately 1 to 100 nm, and the pore volume fraction is less than 30% of the carbon support volume. These micro and meso-pores may be created by chemical etching or chemical reaction of the dense carbon support. One example of chemical etching is plasmas treatment in which the micro and meso-pores are created on the carbon surface by removal of the carbon with oxidation reaction. One example of the chemical reaction is reaction of the carbon with hot KOH in non-oxidizing environment as described by the following equation:

The dense carbon support surface is in contact with molten KOH over a temperature range of 300 to 700° C. in a non-oxidizing environment. The carbon is removed as carbonates, carbon monoxide, carbon dioxide and hydrocarbons. After the reaction, reacted products, such as metallic potassium, carbonates, and hydroxides can be removed by dissolving them in de-ionized water to obtain micro and meso-pores. The pore penetration depth, pore size, and pore volume fraction may be controlled by using proper reaction conditions.

The functionality of the carbon support with micro and meso-pores is illustrated in FIG. 6B. First, those small pores help immobilization of nano-catalyst particles by mechanical interlocking. Second, those small pores help immobilization of a surface layer of electrolyte solution by capillary condensation and surface adsorption so that continuous hydroxide transport paths can be formed throughout the carbon support.

The mass and heat transfer as well as chemical reaction processes occur in the EC cell of the present invention are delineated in the following sections. Ethanol production is used as an example. The processes are applicable to production of other hydrocarbon products. Reaction in the cathode reaction zone to form ethanol:

Oxidation of hydroxide in the anode reaction zone to form oxygen gas:

Minimization of transport resistance of reactants. The reactor productivity is characterized by the molar production rate of product molecule per unit surface area of the membrane, mol/(m²·s), and determined by mass transfer rate and electrocatalytic reaction rate. The mass transfer rate of reactant CO₂ from the cathode channel to the catalytic reaction site, which may be denoted as gas/liquid/solid/electron (G/L/S/e) interface site, may be described by Fick's equation:

J i , m = - D eff , i ⁢ ε c ⁢ C i , f - C i , c δ c , i ( 9 )

where J_i,m=diffusional mass transfer flux of reactant i from feed gas channel to the G/L/S/e interface site in the cathode reaction zone, mol/(m²·s); D_{eff, i}=effective diffusivity of reactant i in the cathode reaction zone, m²/s; ε_c=void fraction of the cathode reaction zone; C_i,f=molar concentration of reactant i in the feed gas, mol/m³; C_i,c=molar concentration of reactant i in the G/L/S/e reaction site, mol/m³; δ_c,i=characteristic mass transfer dimension between the feed channel and cathode reaction zone, m.

The large difference in CO₂ diffusion rate between gas and liquid phase may be shown by ratio of CO₂ gas under atmospheric pressure relative to dissolved CO₂ in liquid water. The Diffusivity of CO₂ is 1.6×10⁻⁵ m²/s in gas phase and 1.6×10⁻⁹ m²/s in water. CO₂ solubility in water @50° C. and 1 bar is 9.4×10⁻³ mol/liter. Ratio of CO₂ gas diffusion rate to its diffusion rate in water is 6,650, that is, the gas diffusion rate is nearly four orders of magnitude of that in water (liquid) phase. Thus, the cathode reaction zone of present invention contains gaseous pores or channels for rapid diffusion of CO₂ reactant from the feed gas to the reaction site.

Electron and ionic transport is the second important design consideration. The minimum voltage required to drive an electrochemical reaction may be calculated from standard Gibbs free energy of the reaction:

E 0 = Δ ⁢ G f n e ⁢ F ( 10 )

Where E₀=thermodynamic minimum voltage to make the electrochemical reaction occur, V; ΔG_f=change of standard Gibbs free energy between the products and reactants, kJ/mol; ne=number of electrons needed to complete the chemical reaction; F=Faraday constant, 96,000 C/mol. E_o is about 1.13V for ethanol production from CO₂ and H₂O.

In an actual cell, there is always electron and electrolyte transport resistance. As a result, actual voltage (E) is higher than the theoretical minimum. The electron transport resistance may be divided into two components: cathode reaction zone, and anode reaction zone. Given current flow density, over voltage may be calculated using the following equation:

Δ ⁢ E e = J e · ( ASR 1 ⁢ e + A ⁢ S ⁢ R 3 ⁢ e ) ( 11 )

Where ΔE_e=over voltage for electron transport, V; J_e=current density, A/cm²; ASR_1e and ASR_3e are areal specific electrical resistance for respective cathode and anode, Ω cm².

To minimize electrical resistance, the cathode and anode reaction zone must have sufficiently high electrical conductivity or sufficiently low electrical resistance. The electrical resistance is preferably less than about 2 Ωcm².

Transport of alkaline ions in the electrolyte between the cathode and anode, such as OH-ions, also imposes over voltage. The ionic transport resistance may be divided into three components: cathode reaction zone, membrane separator, and anode reaction zone. Given current flow density, over voltage may be calculated using the following equation:

Δ ⁢ E ion = J e · ( ASR 1 ⁢ ion + A ⁢ S ⁢ R 2 ⁢ ion + A ⁢ S ⁢ R 3 ⁢ ion ) ( 12 )

Where ΔE_ion=over voltage, V; J_e=current density, A/cm²; ASR_1ion, ASR_2ion, and ASR_3ion are areal specific ionic transport resistance in respective cathode, membrane, and anode, Ω cm².

Ionic transport resistance across the membrane is decreased by use of the thin membrane of pores that are fully wettable by electrolyte solution. Ionic resistance in the anode is minimized by keeping anode in the liquid electrolyte solution. The ionic resistance in the cathode zone is reduced by 1) reducing thickness of the catalyst coating/layer, and 2) having the ionic conducting paths in the catalyst coating/layer. The ionic conducting paths may be formed by capillary condensation of the liquid electrolyte among interparticle and intraparticle voids and/or by adsorption of the liquid electrolyte on the catalyst and/or support surface. The ionic transport resistance through the membrane separator is reduced by using a thin membrane. The membrane thickness is preferably less than 0.2 mm, preferably less than 0.05 mm.

Cell temperature control. The over-voltage will be converted to Joule heating inside the cell. The heating power may be described by the following equation:

q e = J e · ( Δ ⁢ E e + Δ ⁢ E i ⁢ o ⁢ n ) = J e · ( E - E 0 ) = J e · E · ( 1 - η e ⁢ c ) ( 13 )

Where q_e=joule heating power density, w/cm²; J_e=current density, A/cm²; E=operating voltage, V; η_EC=electrical efficiency of the electrochemical cell.

The electrochemical cell needs to be operated at current density high enough to achieve high productivity, as expressed by the following equation:

J P = γ P ⁢ J e n e ⁢ F ( 14 )

Where J_P=productivity of product molecules, mol/(m²·s); γp=product selectivity, fraction of total electrons used toward producing product molecule.

q l = q e ( 15 ) q l = J 2 ⁢ l · C p , l · Δ ⁢ T ( 16 )

Where q_l=heat removal rate per unit cell area through liquid flow through the anolyte flow channel (zone 2), W/m²; q_e=Joule heating flux, W/m²; J_2l=liquid electrolyte flow rate in the anolyte flow channel per unit cell area, Kg/(s·m²); C_p,l=specific heat capacity of the liquid fluid, KJ/Kg/K; ΔT=the liquid fluid temperature increase, K.

Temperature control of the cell is illustrated with four cases for electrochemical production of ethanol. The results are summarized in Table 1. Case 2 has higher electrical efficiency (η_ec). As expected, Joule heating flux is significantly reduced from 1370 W/m² at 45% efficiency to 870 W/m² at 57%, while the current density and productivity are kept the same. In cases 2,3, and 4, Joule heating flux proportionally increases with current density under constant electrical efficiency. Large amounts of heat are produced when the cell is operated at high productivity. Effective removal of such heats is essential to operate the cell within certain range of temperature. The cell operating temperature in the present invention is preferably between about 20 to about 95° C., i.e., between room temperature and water boiling temperature. More preferably, the temperature is above 35° C. so that no refrigeration cooling is needed in warm climates.

As shown in FIGS. 2A and 2B, the Joule heat is removed by introducing liquid-phase electrolyte fluid into the anode flow channel. The liquid fluid may be aqueous carbonate, bicarbonate, hydroxide, or their mixtures of pH greater than 7. The anode channel is characterized by channel spacing (or gap), channel span, and cross-sectional flow area. In instances in which the channel spacing is 1 mm and the channel span is 0.5 m per m² of the cell, specific channel flow area is 0.0005 m²/m². At the liquid flow rate of 0.1 kg/s per m² of the cell area, the liquid fluid temperature increase through the cell may be ranged from 2.2 to 8.7K. The corresponding mass ratio of the liquid fluid to the product ranged from about 700 to 3000. In other words, mass flow rate of liquid fluids needs to be two to three orders of magnitude higher than the product formation rate.

Heat removal by gas flow in the cathode flow channel (6) is much less effective than that by liquid flow. The exemplary gas, as shown in table 1 (below), has density and specific heat capacity less than the liquid by about 500× and 4×, respectively. At the same mass flow rate as the liquid, the gas temperature increase through the cell is about 4× of the liquid fluid. However, such high gas mass flow rate would result in exceedingly high gas velocity and pressure drop, which could substantially increase the cell fabrication cost. If the specific area for gas flow in the cathode channel is same as the anode channel, the gas velocity could need to be as high as 115 m/s for removal of Joule heat, which is too high for practical devices.

The cell design of present invention uses the liquid fluid in the anode channel for removal of Joule heat, while the cathode channel is for introduction of feed gas and discharge of gaseous products. In instances in which feed gas mass flow is ten times of the product formation rate, the gas velocity is in the range of 0.3 to 1.5 m/s, which is within the practical operation range.

The cell is designed with adequate thermal conductivity so that the Joule heat may be rapidly transported to the anode channel. For illustration purposes, the cathode reaction zone/membrane/anode reaction zone is treated as one whole heat generation body. Thermal conduction is the primary heat transfer mechanism as described below:

J q = k eff ⁢ Δ ⁢ T δ eff ( 17 )

In instances in which the effective thermal conductivity (k_eff) for the cell body is 2 W/(m·K) and characteristic heat transfer dimension (δ_eff) from the cell body to the anode channel is 2 mm, Table 1 shows that the temperature gradient to transport of the Joule heat is in the range of 1 to 4K.

For comparison purposes, thermal conductivity (W/(m·K)) is 0.028 for air, 0.6 for water, 15 for stainless steel, 0.6 for activated carbon, 100 for nickel, and 400 for copper. Thermal conductivity (W/(m·K)) of carbon-type materials varies over a very broad range, from 0.1 to 0.6 for amorphous-type carbon, 200 to 800 for graphite flakes, 24 W/(m·K) for basic graphite fibers, 1200 for high-modulus pitch-based fibers, from 2000 to over 5000 for graphene, to 6,600 for high-quality, individual single-walled carbon nanotubes. Graphite particles, flakes and fibers are available economically, and thus, may be used to reduce fabrication cost of the cathode reaction zone.

TABLE 1
Joule heating and temperature control of the cell for production of ethanol

Case	1	2	3	4

Working voltage	2.5	2.0	2.0	2.0
Current density, mA/cm2	100	100	200	400
Power consumption, W/m2	2,500	2,000	4,000	8,000
Electrical efficiency	45%	57%	57%	57%
Joule heating flux, W/m2	1,370	870	1,740	3,480
Selectivity to ethanol	0.80	0.80	0.80	0.80
Overall efficiency to C2 product	0.36	0.45	0.45	0.45
Cell productivity,
mol/(m2 · s)	6.9E−04	6.9E−04	1.4E−03	2.8E−03
kg/(m2 · s)	3.2E−05	3.2E−05	6.4E−05	1.3E−04
Removal of Joule heat by liquid
Specific channel flow area, m2/m2	5.0E−04	5.0E−04	5.0E−04	5.0E−04
Anolyte liquid flow, kg/(m2 · s)	0.10	0.10	0.10	0.10
Density, kg/m3	1,000	1,000	1,000	1,000
Specific heat capacity, kJ/(kg · K)	4.00	4.00	4.00	4.00
Mass ratio of liquid/product, kg/kg	3,130	3,130	1,565	783
Liquid velocity, m/s	0.20	0.20	0.20	0.20
ΔT, K	3.4	2.2	4.4	8.7
Removal of Joule heat by gas
Specific channel flow area, m2/m2	5.0E−04	5.0E−04	5.0E−04	5.0E−04
Cathode gas flow, kg/(m2 · s)	0.10	0.10	0.10	0.10
Specific heat capacity, kJ/(kg · K)	1.00	1.00	1.00	1.00
Mass ratio of liquid/product,	3,130	3,130	1,565	783
Kg/Kg
Gas velocity, m/s	115	115	115	115
ΔT, K	13.7	8.7	17.4	34.8
Thermal conduction inside the cell
Effective thermal conductivity,	2.0	2.0	2.0	2.0
w/(m · s)
Characteristic heat transfer	2.0E−03	2.0E−03	2.0E−03	2.0E−03
dimension, m
ΔT, K	1.4	0.9	1.7	3.5

Product mass transfer. In addition to feed CO₂ mass transfer and heat transfer discussed above, mass transfer of product molecules is the other cell design consideration. These common C1, C2, and C3 molecules as listed in Table 2 have a broad range of boiling points. Except for ethylene, all these oxygenated molecules have boiling points above room temperature (20-25° C.) at which most electrochemical cells reported in the literature operate. In other words, all those oxygenated molecules are in liquid phase if the cell is operated at room temperatures or at temperatures below the boiling point. Mass transfer diffusion rate in liquid phase is much less in the gas phase. The product mass transfer process is described with ethanol product as an example. The approaches are applicable to other molecules.

TABLE 2
Boiling points of a list of molecules with small carbon number

	Name	Boiling point, ° C.

	C1	Methanol	65.0
	C1	formic acid	100.8
	C2	ethanol	78.0
	C2	ethylene	−103.7
	C2	acetic acid	117.9
	C3	propanol	97.0
	C3	isopropanol	80.4
	C3	acetone	56.5

As shown in FIGS. 2A and 2B, a product molecule formed in the cathode reaction zone may diffuse into the cathode channel by gas diffusion and into the anode channel by liquid diffusion.

Diffusion rate of product i from the cathode reaction zone into the cathode channel is described by the following equation:

J i , 56 = D i , 56 ⁢ C i , 5 - ⁢ C i , 6 δ 5 ⁢ 6 ( 18 )

Assume idea gas

C i , 6 = y i , 6 · p 1 R ⁢ T · 10 3 ( 19 )

Assume that gas pressure in the cathode channel is same as in the cathode reaction zone. If product molecule i is in vapor phase in the cathode reaction zone at reaction temperature T,

C i , 5 = y i , 5 · p 1 R ⁢ T · 10 3 ( 20 )

In instances in which a product molecule i is in liquid phase at reaction temperature T, its gaseous concentration may be approximately described by

C i , 5 = p i , eq · x i , 5 R ⁢ T · 10 3 . ( 21 )

Where J_i,56 is diffusional mass transfer rate of product i from the cathode reaction zone (5) to the cathode channel (6), mol/(m²·s); D_i,56 is effective diffusivity of product i from the cathode reaction zone to the cathode channel m²/s; C_i,5 is molar concentration of product i in the cathode zone, mol/m³; C_i,6 is molar concentration of product i in the cathode channel, mol/m³; δ₅₆ is characteristic mass transfer dimension from the cathode reaction zone (5) to the cathode channel (6), m; p₁=gas pressure in the cathode channel, bar; y_i,6 is molar fraction of product i in the cathode channel; p_i,eq=equilibrium pressure of product i in the electrolyte solution at reaction temperature; x_i,5 is molar fraction of product i in the electrolyte solution in the cathode zone; y_i,5 is molar fraction of product i in vapor phase on in the cathode zone.

Liquid-phase diffusion rate of product i from the cathode reaction zone (5) into the anode channel (2) is described by the following equation:

J i , 52 = D i , 52 ⁢ C i , 5 - ⁢ C i , 2 δ 52 ( 22 )

Where J_i,52 is diffusional mass transfer rate of product i from liquid electrolyte in the cathode reaction zone (5) to the liquid fluid in the anode channel (2), mol/(m²·s); D_i,52 is effective diffusivity of product i in the liquid electrolyte from the cathode reaction zone to the anode channel m²/s; C_i,5 is molar concentration of product i in the electrolyte solution in the cathode zone, mol/m³; C_i,2 is molar concentration of product i in the liquid fluid in the anode channel, mol/m³; δ₅₂ is characteristic mass transfer dimension from the cathode reaction zone (5) to the anode channel (2).

Under steady-state operation, the diffusional mass transfer rate is equal to the product formation rate:

J i , p = J i , 56 + J i , 52 ( 23 )

Product mass transfer under different operation conditions is calculated with the above equations to show design features of the cell design. The results are summarized in Table 3. The liquid product may come out of the cell from both cathode and anode effluents. Comparison of cases 1 to 2 shows impact of the reaction temperature on product distribution when ethanol concentration in the anode flow channel is controlled at 0.001 M. Comparison of cases 3 to 4 shows impact of the reaction temperature on product distribution when ethanol concentration in the anode flow channel is controlled at 0.5 M. Comparison of cases 4 to 5 shows impact of the changing ethanol concentration in the anode channel on product distribution at a constant reaction temperature. To increase product content in the cathode and anode effluent, the cell may be operated at elevated temperatures. The cell may be operated at low pressures to keep cathode channel at vapor phase and facilitate diffusional mass transfer.

The product concentration in the electrolyte could be high. Oxygenated hydrocarbons such as alcohol are strong solvents for dissolution of organic materials. Thus, the cell is preferably made of materials that are durable in the presence of hydrocarbon liquid and/or gas at elevated temperatures.

TABLE 3
Product mass transfer in the electrochemical cell under different
reaction temperatures and ethanol concentration in the anode channel
while product productivity and reactor pressure kept constant.

Case #	1	2	3	4	5

Cell operation conditions

Product formation	6.9E−04	6.9E−04	6.9E−04	6.9E−04	6.9E−04
rate, mol/(m2 · s)
Pressure in gas	1.1	1.1	1.1	1.1	1.1
channel, bar
Temperature, ° C.	20	50	50	78	78
Ethanol conc. in	0.001	0.001	0.500	0.500	2.000
liquid-phase anode
channel, M

Ethanol in the cathode reaction zone

Peq of ethanol, bar	0.07	0.25	0.25	1.00	1.00
Ethanol conc. in	2.20	1.07	1.56	0.88	3.23
electrolyte, M
Partial pressure of	0.003	0.005	0.007	0.016	0.058
ethanol vapor, bar

Vapor-phase mass transfer from the cathode reaction zone to cathode channel

Void fraction in cathode	0.20	0.20	0.20	0.20	0.20
Cathode mass	0.001	0.001	0.001	0.010	0.100
transfer dimension, m
Ethanol molar fraction	0.001	0.001	0.003	0.010	0.050
in channel gas
Effective gas	2.6E−06	3.1E−06	3.1E−06	3.6E−06	3.6E−06
diffusivity, m2/s
Gas transfer	23.1%	63.3%	63.4%	87.1%	57.6%
rate/formation rate

Liquid-phase mass transfer from the cathode reaction zone to anode channel

Anodic mass	0.001	0.001	0.001	0.001	0.001
transfer dimension, m
Void fraction	0.20	0.20	0.20	0.20	0.20
Effective	2.4E−10	2.4E−10	2.4E−10	2.4E−10	2.4E−10
diffusivity, m2/s
Liquid transfer	76%	37%	37%	13%	43%
rate/formation rate

Electrochemical conversion process. A gas stream containing CO₂ and H₂O is introduced into cathode gas inlet of the EC device (FIG. 1) at process conditions of pressure=P_g,in, temperature=T_g,in, and flow rate=F_g,in. The un-converted gas and vapor-phase products come out of the device from cathode gas outlet at process conditions of pressure=P_g,out, temperature=T_g,out, and flow rate=F_g,out. An electrolyte liquid flow of pH>7, may be also called as anolyte, is introduced into the anodic liquid inlet at process conditions of pressure=P_l,in, temperature=T_l,in, and flow rate=F_l,in. The anolyte and reaction products come out of the anodic liquid outlet at process conditions of pressure=P_l,out, temperature=T_l,out, and flow rate=F_l,out. An electrical power is applied to the cathode side and anode side of the device at voltage less than 3.0V for each EC cell and with current density of greater than 50 mA per cm² of the membrane working area. In an EC cell, CO₂ and H₂O in the cathode flow channel diffuse into the catalytic reaction site in the cathode reaction zone and are converted into hydrocarbon products and hydroxide ions by taking electrons from the cathode reaction zone. The vapor-phase hydrocarbon product diffuses into the cathode flow channel from the catalytic reaction site, while the liquid-phase hydrocarbon products diffuse across the membrane separator into the anode flow channel. The hydrocarbon products include methanol, ethylene, ethanol, acetate, propanol, acetone, and their mixture. The hydroxide ions diffuse across the membrane separator into the anode reaction zone where the hydroxide ions are converted into oxygen gas and water by releasing electrons to the anodic reaction zone. Oxygen gas products diffuse from the anodic reaction site into the anode flow channel. The cathode gas pressure is always maintained higher than the anode liquid pressure to prevent crossover of the gas and liquid, that is, P_g,in or P_g,out>P_l,in or P_l,out.

The cathode gas inlet pressure and temperature are preferably controlled in the ranges of about 1.1 to about 3.0 bar (absolute) and 35 to 95° C., respectively. The cathode gas pressure drop from inlet to outlet is preferably controlled less than 0.5 bar. The cathode gas temperature increase from inlet to outlet is preferably controlled less than 50° C. CO₂ content in the cathode feed gas is preferably greater than 50 vol %. Oxygen content in the cathode feed gas is preferably less than 5 vol %. The anode liquid inlet pressure and temperature is preferably controlled in the ranges of about 1.0 to about 2.0 bar (absolute) and 20 to 80° C., respectively. The anode liquid flow pressure drop from inlet to outlet is preferably controlled less than 0.5 bar. The anode liquid temperature increase from inlet to outlet is preferably controlled less than 20° C. The anode liquid comprises a solution of carbonates, bicarbonates, hydroxides, and their mixture with molar concentration of 0.1 to 5M and PH>7.

Example I. Structure and Performance Characteristics of Zirconia/Nickel Sheet

Zirconia is known as a durable membrane material. Yttria-stabilized zirconia powder of average crystalline sizes of about 200 nm is impregnated with NiO sintering promoter. The powder is attrition milled to form a stable coating solution with addition of dispersants and binders in an organic solvent. 210 mm×210 mm porous Nickel sheet is coated with the coating solution under vacuum filtration. The coated sheet is dried under environmental conditions without any cracks. Then, the coated sheet is cut into 10×4 cm×10 cm coupons. The coupons are loaded inside a tubular reactor, purged with nitrogen gas, and sintered in hydrogen gas flow by gradually raising temperature to 750-800° C. and holding for 2 hours.

Two groups of ceramic/Nickel sheets are prepared using porous Nickle sheets of different Nickel content. The porous Ni sheet has thickness ranged of 71 to 78 μm, porosity ranged from 45 to 48%. Table 4 lists the first group of sheets coated on the porous Nickel sheet of about 97 at. % Ni. The coupons without short circuits and defects are used for cell assembly. In general, all the five sheets show good adhesion of the ceramic coating on the Nickel sheet. For the first three sheets, little delamination and wrinkles are observed with only two coupons, while all the coupons for the latter two sheets show no delamination, no cracks, and no deformation.

Table 5 lists the second group of zirconia coatings on the nickel sheet containing 99 at % Nickel. There is serious delamination with the first three sheets, while the latter two sheets show no delamination, no cracks. Thicknesses of the zirconia coating on the first three sheets are ranged from about 26 to about 57 μm, while the coating thicknesses for the latter two sheets are about 15 to about 24 μm. The results confirm that the ceramic coating thickness is preferred to be less than about 25 μm. Taking into account the results in Table 4, the coating thickness is preferred to be greater than about 9 μm.

TABLE 4
Zirconia coating on porous Nickel alloy sheet of 97%
(210 mm × 210 mm sheet size, one-layer coating)

Ni alloy material	BNO	BNO	BNO	BNO	BNO

Ni sheet ID	5603	5599	7705	7693	7692
Avg thickness,	78	76	73	74	71
um
Porosity	0.47	0.47	0.49	0.45	0.45
Ceramic
coating
Solid in sol,	1.71	1.71	1.71	1.1	1.1
wt. %
Volume used, ml	120	160	200	200	200
Areal loading	6.2	8.5	10.9	5.0	5.0
density, mg/cm2
Observation of	little	little	Little	No	No
after sintering	delamination/	on 2 coupons	on 2 coupons	or wrinkles,	delamination
	cracks/wrinkles	delamination	delamination	shiny, no cracks.	or wrinkles,
	on 2 coupons			delamination	no cracks,
					shiny
Sintered
coupon A
Avg thickness,	85	105	104	83	86
um
STDEV, um	2	2	1	1	2
Coating	7	29	31	9	15
thickness, um
Sintered
coupon B
Avg thickness,	100	98	105	83	80
um
STDEV, um	3	1	7	1	1
Coating	22	22	32	9	9
thickness, um

TABLE 5
Zirconia coating on porous Nickel alloy sheet of 99%
(210 mm × 210 mm sheet size, one-layer coating)

Ni alloy material	GNO	GNO	GNO	GNO	GNO

Ni sheet ID	6499	6496	7243	7254	7248
Avg thickness, um	79	79	99	95	97
Porosity	0.45	0.45	0.45	0.45	0.45
Ceramic coating
Solid in sol, wt. %	1.71	1.71	1.71	1.1	1.1
Volume used, ml	120	160	200	200	200
Areal loading	5.8	8.5	11.4	5.8	6.2
density, mg/cm2
Observation of	Delaminated	Delaminated/	Small	No	No
after sintering	and wrinkled	wrinkled	delamination	delamination,	wrinkles
			on edge	no cracks	or cracks,
					uniform
					coating.
Sintered
coupon A
Avg thickness, um	113	131	141	119	113
STDEV, um	4	5	5	8	4
Coating thickness,	34	52	42	24	16
um
Sintered
coupon B
Avg thickness, um	105	136	142	110	113
STDEV, um	1	7	2	2	3
Coating thickness,	26	57	43	15	16
um

Representative structures of zirconia/Nickel sheets are shown in FIGS. 7A-7C. The fractured wall (FIG. 7A) reveals the zirconia coating thickness of about 10 μm and penetration into underneath pores at a depth of about 10 μm, and porous structures of the Nickel sheet throughout its whole thickness. FIG. 7B shows uniform pores on the Nickel sheet surface at micro and sub micrometer sizes. FIG. 7C shows that the Nickel sheet surface is fully covered by the zirconia coating. The zirconia coating has much smaller pores than the Nickel sheet and no cracks.

The gas permeation characteristics of the zirconia/Nickel sheet is evaluated by measuring permeation rate of humid air at different pressure gradients. The results are plotted in FIG. 8. With as-sintered sheet, air permeation rate at 22% and 76% relative humidity increases with pressure gradient. The membrane is permeable to air at almost zero pressure gradient. After the zirconia membrane pore is filled with KOH electrolyte solution, the membrane has a small air permeation rate with air of 44% humidity until a breakthrough pressure (˜8 kPa). Beyond the breakthrough pressure, the air permeation rate increases with pressure gradient. With 78% RH air, no air permeation could be measured at the pressure gradient evaluated, which is up to 14 kPa for this set of testing. The results confirm that the pores are fully blocked by the condensed water and electrolyte solution.

In the EC cell of the present invention, the porous Nickel side of the zirconia-coated Nickel sheet is exposed to the anode electrolyte liquid to assure that the zirconia membrane pores be filled with the electrolyte solution all the time under the reaction conditions.

Activity of the porous Nickel alloy sheet for water electrolysis is assessed by measuring water electrolysis current at different voltages. A test cell is assembled by clamping a porous Nickel sheet with zirconia/Nickel sheet. The cell assembly is immersed in an electrolyte solution at room temperature. The water electrolysis activity is compared in FIGS. 9A and 9B with 1.0 M KOH and 1.0 M KHCO; electrolyte solutions. The testing results plotted in FIG. 9A are obtained by using the Nickel sheet as positive electrode and zirconia-coated Nickel sheet as negative electrode. In FIG. 9B, the Nickel sheet and zirconia-coated Nickel sheet are used as respective negative and positive electrode. As expected, the porous Nickel sheet is active for electrolysis of KOH solution as either cathode or anode. The Nickel sheet shows little water electrolysis activity in the KHCO₃ solution at ≤2.0V. Thus, using KHCO₃ solution in the present EC cell may minimize water electrolysis reaction.

Example II. Electrochemical Conversion of CO₂ and H₂O into C₂₊ Oxygenated Fuels

Feasibility of the proposed reactor design and process method is demonstrated on laboratory apparatus with process flow diagram in FIG. 10. CO₂ gas produced from direct air capture (DAC) regeneration process is stored in vessel C21 by pushing water to another vessel C22 located at a higher position. The stored CO₂ gas is introduced to cathode side of an electrochemical test cell (EC-1) through an in-situ humidifier (HD1). The EC cell and humidifier are placed inside an oven for temperature control. The cathode effluent gas is cooled at room temperature, and the condensate is collected in a sample tube (KD1). The remaining gas is filtered through a membrane filter and ethanol content in the exhaust gas is sampled by a portable ethanol analyzer. 1.0 M potassium bicarbonate solution is used as the electrolyte and circulated on the anode side of the EC1. The gas coming out of the anode side is separated from the electrolyte solution in a beaker (KD2). Thermometers are placed on the cathode and anode outlets of the EC cell, respectively. A manometer is used to monitor pressure difference between the cathode and anode.

The EC1 cell is assembled using two stainless steel cover plates (FIG. 11A). The cover plate has inlet and out ports for connection to external tubes. Flow channels with 2.0 mm width×1.0 mm depth are carved on interior side of the cover plate. The inlet and outlet ports on exterior of the cover plate are connected to beginning and ending of the flow channel, respectively. Thus, the end plate and flow channel are built as one piece of plate in this experimental device. The zirconia/Nickel sheet as shown in example I is used as anodic membrane electrode assembly (MEA), i.e, the ceramic membrane separator is integrated with the anode electrode (anode reaction zone) into one piece. The membrane sheet used for this test cell is 3.0 cm width×5.0 cm length. The edge of the membrane at about 5 mm-wide is covered by a PTFE film to facilitate sealing of the cathode from anode side. Thus, the active working area of the test cell is 2×4=8 cm². The cathode reaction zone is made as a copper-coated porous carbon paper. The copper catalyst coating thickness is about 300 to 600 nm. The carbon paper thickness is about 0.3 mm. The cathode cover plate/carbon paper-copper coating/ceramic coating-porous Nickel sheet/anode cover plate are stacked together using an EPDM gasket and compressed uniformly by threaded screws to obtain a working cell. The uniform compression is necessary to seal the cathode and anode flow channels from environment by the EPDM gasket and is also necessary to obtain intimate contact with all the components. The compression force is properly controlled to ensure the copper catalyst coating in intimate contact with the ceramic membrane surface without causing breakage of the carbon paper and/or the ceramic coating. Both carbon paper and ceramic coating are fragile.

The assembled EC cell is placed inside an oven for testing. The cathode and anode end (or cover) plates are connected to respective negative and positive ports of a DC power supply. The dry cell is tested by applying voltages in the range of 0.5 to 2.5V to assure no current. CO₂ gas is first introduced into the cathode flow channel at a flow rate of 0.1 liter/min to build a positive pressure. Then, liquid electrolyte flow is introduced into the anode flow channel at a flow rate of 17.5 cc/min. After the testing cell is checked without gas leakage, liquid leakage, and gas/liquid crossover, then oven temperature is increased to 50-54° C. The electrochemical reaction starts in response to electrical power being applied to the cell. Testing results are summarized in Table 6.

At 1.5 V, current density was 65 mA/cm². When the voltage was raised to 2.0, 2.5, and 3.0V, the current density rapidly increased to 193, 330, and 366 mA/cm², respectively. Throughout the whole reaction testing period, the liquid electrolyte temperature at the cell outlet was about 38 to 41.8° C. Since the electrolyte reservoir was kept at room temperature (˜25° C.), the electrolyte temperature increase was due to heat transfer from the cell to the liquid. Since heat transfer rate from oven to the electrolyte solution is relatively slow, the electrolyte temperature at the cell outlet was less than the oven temperature. By contrast, the gas temperature at the cell outlet was higher than the oven temperature. As the voltage was raised, the gas outlet temperature became higher and higher than the oven temperature. The gas temperature increase suggests that in this cell assembly, heat transfer rate from the cathode reaction zone to the anode flow channel was not rapid enough relative to the heat generation rate.

TABLE 6
Electrochemical reaction activity and temperature control with cell assembly 1

				Cathode	Anode
				feed gas	electrolyte	Electrolyte	Gas T at
			Over	rate,	circulation	T at cell	cell outlet
Time	V	mA/cm2	T, ° C.	L/min	rate, cc/min	outlet, ° C.	T, ° C.

14:00	startup
14:20	1.5	65	53.8	0.1	17.5	38.1	54
14:25	1.5	65	50.8	0.1	17.5	38.2	54
14:26	2.0	178	49.9	0.1	17.5	38.4	58
14:29	2.0	190	53.7	0.1	17.5	39.5	65
14:34	2.0	193	51.8	0.1	17.5	39.7	67
		0
14:37	2.5	323	53.1	0.1	17.5	40.4	76
14:42	2.5	324	51.4	0.1	17.5	40.8	86
14:43	2.5	330	53.7	0.1	17.5	40.9	87
		0
14:44	3.0	411	53.8	0.1	17.5	41.6	99
14:47	3.0	395	50.3	0.1	17.5	41.7	114
14:49	3.0	378	50.4	0.1	17.5	41.6	115
14:51:00	3.0	366	54.0	0.1	17.5	41.8	116

Another cell was assembled with the components and procedure as described above. The testing results are summarized in Table 7. With cell assembly 2, heat transfer from the cathode reaction zone to the anode flow channel was substantially enhanced. Gas temperature at the cell cathode side outlet was about same as the electrolyte at the cell anode side outlet, during the whole testing process. It is noted that when the CO₂ feed gas was turned off in the middle of the run, current density dropped from 165 to 26 mA/cm² in 1 min. After the CO₂ feed gas was resumed, the current density rapidly recovered to 151 mA/cm². This testing run confirms that the electrochemical reaction was dominated by CO₂ reduction.

TABLE 7
Electrochemical reaction activity and temperature control with cell assembly 2

				Cathode	Anode
				feed gas	electrolyte	Electrolyte
			Over	rate,	circulation	T at cell	Gas T at cell
Time	V	mA/cm2	T, ° C.	L/min	rate, cc/min	outlet, ° C.	outlet T, ° C.

16:39			52	0.08	0		29.1
16:42	2.00	151	49.9	0.1	5	48.7	48.3
16:47	2.00	165	52.3	0.1	5	19.3	49.6
17:01	2.00	165	53.6	0.1	5	49.2	49.8
17:08	2.00	165	53.8	0.1	5	48.9	49.5
17:09	2.00	26	53	0	5	49.1	49.7
17:12	2.00	151	51.1	0.06	5	48.7	48.6
17:19	2.00	159	51.6	0.06	5	48.9	49.1

Third cell was assembled and tested with the method as described above. The cell was tested with CO₂ feed gas flow rate of 0.1 liter/min and liquid electrolyte solution circulation rate of 5 cc/min at oven temperature of 50° C. The cathode effluent gas was condensed at room temperature to obtain liquid samples, while the liquid electrolyte was sampled from the anode effluent. The sample was analyzed by NMR technique. The results are summarized in Table 8. Four pairs of the samples were obtained at different voltages. FIG. 12 shows representative NMR peaks for a cathode condensate sample (#7). DMSO here is used as internal standard for quantitative analysis. CH₃OH, acetone, acetate, ethanol, and isopropanol (IPA) are identified as the major constituents. A few hydrocarbon peaks could not be identified.

The analysis shows presence of reaction products in both cathode condensate and anode electrolytes, as expected based on the above product mass transfer discussion. C2 and C3 oxygenates are the major products. The electrochemical reaction also produced formate and methanol, these C1 oxygenates. It is worth noting that C2 and C3 product concentrations did not increase with voltage. Joule heating increases with voltage. Thus, the cell design and assembly should be optimized to operate at voltages at 2.5V or below.

TABLE 8
Molar concentrations of the cathode effluent condensate and anodic electrolyte

#	Effluent	Date	Voltage, V	Formate	Methanol	Ethanol	Acetate	Isopropanol	Acetone

3	Cathode	Day 1	2.0	2.4E−04	1.9E−03	1.7E−03	2.6E−04	7.0E−03	4.1E−04
2	Anode	Day 1	2.0	0.0E+00	1.8E−05	0.0E+00	6.5E−05	3.3E−03	5.1E−05
5	Cathode	Day 1	2.5	1.6E−04	3.7E−04	5.4E−04	2.1E−04	2.8E−03	9.7E−04
4	Anode	Day 1	2.5	0.0E+00	1.7E−05	0.0E+00	1.0E−04	3.0E−03	2.7E−04
7	Cathode	Day 2	3.0	1.8E−04	6.7E−04	6.1E−04	2.9E−04	1.5E−03	1.1E−03
6	Anode	Day 2	3.0	0.0E+00	2.6E−05	0.0E+00	1.6E−04	2.4E−03	4.6E−04
9	Cathode	Day 2	3.3	1.1E−04	9.1E−05	1.7E−04	1.6E−04	3.4E−04	4.9E−04
8	Anode	Day 2	3.3	0.0E+00	2.0E−05	0.0E+00	1.9E−04	2.0E−03	5.8E−04

Example III. Stability of Electrochemical Conversion of CO₂ and H₂O

Cell assembly 4 was made using carbon sheet-supported copper catalyst (denoted as Cu/BP-C) with the configuration as shown in FIGS. 11A and 11B, except that a 99.9% pure copper gauze mesh 100 was placed between the cathode cover plate and the Cu/BP-C cathode catalyst sheet. The testing setup was the same as shown in FIG. 10 except that CO₂ from beverage grade CO₂ cylinder was used to provide stable and consistent CO₂ gas for stability assessment. Ethanol content in the cathode exhaust gas was assessed using an ethanol breath analyzer. The analyzer was calibrated by bubbling CO₂ gas into ethanol/water inside a sample tube with the same gas flow rate and sample tube size. FIG. 13 shows correlation of the breather analyzer reading with ethanol content in the liquid being purged by gas. The ethanol analyzer has an upper detection limit of 19.0.

It was noticed that the new cell assembly required a certain period for activation. To shorten the activation time, testing of the cell assembly started at 3.0V. Results in Table 9 show that the cell was activated within about 1 h. A positive pressure gradient between the cathode and anode was maintained by using the disk membrane filter of suitable pore size in the cathode effluent downstream. The cathode affluent coming out of the filter was analyzed by a portable breath analyzer for qualitative assessment of presence of ethanol. During about 80-min run at 3V, ethanol reading of the breath analyzer reached its detection limit, 19.0. Then, the voltage was reduced to 2.5V, and the cell was left to run overnight at current density of 122 mA/cm². In the next day morning, the cell current density stabilized at 172 mA/cm², while ethanol reading remained at the detection limit of 19.0. Through the whole testing process, the cell temperatures at the cathode and anode outlets were well controlled, the positive pressure was maintained between the cathode and anode, no CO₂ gas permeation into the anode occurred, and no permeation of the electrolyte solution into the cathode was observed. After overnight, the pressure gradient increased to 19 kPa from 4-5 kPa, because of partial blocking of the membrane filter in the cathode downstream. It is worth noting that no gas crossover occurred at such a higher-pressure gradient, confirming that the capillary pressure of the present zirconia membrane may withstand a significant pressure gradient between the gas and liquid sides.

Then, the cathode effluent line was disconnected to take the condensate sample from the knock drum (KD1) and clean the membrane filter, and the anode effluent line was disconnected to collect the electrolyte sample. After sampling, the lines were reconnected. The pressure gradient came back to the previous level of about 3-6 kPa. However, the ethanol reading of the cathode exhaust suddenly dropped to 1.1 and 0. It was suspected that the cathode may have been incidentally flooded with the electrolyte solution during the sampling. Thus, the CO₂ feed gas rate was doubled to 0.2 liter/min and the oven temperature increased to 80° C. Within about 1.5 h, the ethanol reading recovered to 19.0 as the oven temperature reached 80° C. At 80° C. oven temperature, the current density gradually increased to 191 mA/cm² in 3 hours while the ethanol reading remained at 19.0. It is noted that the cell cathode side outlet temperature was about 1 to 2° C. higher than the cell anode side outlet. Both the cathode and anode outlet temperatures were less than the oven temperature, which is explained by cooling effects of feed gas and circulating electrolyte at room temperature (22-25° C.).

The testing results of the cell assembly 4 confirmed that there was no deactivation of the catalyst and cell performance with time. Instead, the cell tended to become more active with time. By maintaining the pressure differential between the cathode and anode side, gas and liquid crossover could be avoided. The cathode and anode side temperature could be well controlled. The catalyst and cell may work at temperatures of 50 to 80° C., above the room temperature. The higher the cell reaction temperature is, the less product separation cost could be, and higher the value of the heat rejected by the cell reaction could be to direct air capture regeneration process.

TABLE 9
Stability and re-activation of the cell assembly 4 loaded with cathode catalyst sheet of Cu/BP-C

Ethanol

Sampling

Oven T, ° C.

Cathode

Anode

Fluid T coming

in

Time,

Power

Set

gas

electrolyte

out of the cell, ° C.

cathode

min	V	mA/cm2	point	actual	L/min	cc/min	Electrolyte	Gas	Δp, Pa	exhaust

Day 1 startup

13:03	3.00	165	50	50.0	0.10	5.0	45.2	42.1	5,251	19.0
13:07	3.00	240	50	52.8	0.10	5.0	46.4	43.1	5,539	19.0
13:15	3.00	361	50	51.2	0.10	5.0	47.5	45.1	5,462	19.0
14:22	3.00	398	50	52.2	0.10	5.0	50.3	50.7	4,710	19.0
17:01	2.50	68	50	52.3	0.10	5.0	47.7	47.7	4,477	19.0
17:20	2.50	122	50	50.0	0.10	5.0	48.2	47.9	4,385	19.0

Day 2 after overnight run

9:08	2.50	174	50	53.8	0.10	5.0	49.0	49.9	18,930	19.0
9:20	2.50	172	50	48.8	0.10	5.0	50.0	49.9	NA	19.0

After collecting the gas condensate and electrolyte sample

9:47	2.41	520	50	36.9	0.10	5.0	41.4	40.0	3,131	1.1
11:51	2.05	520	50	53.2	0.10	5.0	52.7	53.4	5,878	0.0

Re-activation of the cell by increasing temperature

12:00	1.50	277	80	79.2	0.20	5.0	67.0	60.1	6,130	0.0
13:35	2.50	165	80	80.0	0.20	5.0	75.0	73.7	OFF	19.0
14:37	2.50	128	80	79.2	0.20	5.0	76.0	74.0	OFF	19.0
15:22	2.50	162	80	83.4	0.20	5.0	76.7	75.3	4,258	13.6

Example IV. Stability of Electrochemical Conversion of CO₂ and H₂O Using a Different Cathode Catalyst Sheet

Cell assembly 5 was made in the same way as cell 4 except for that different type of copper/porous carbon sheet was used, denoted as Cu/GDL-C. The cell was tested in the same way as in example III. Table 10 lists the cell performance at different sampling times under constant oven temperature, feed CO₂ gas flow rate, and electrolyte circulation rate. The pressure gradient between the cathode and anode was maintained at 5 to kPa that no gas or liquid crossover occurred. The cell temperature was controlled well during 2.5-day testing. In the first day startup, the current density increased from 160 to 207 mA/cm² in 40 min and gradually increased to 425 mA/cm² overnight and stabilized at 385 mA/cm². The current density dropped to 283 mA/cm² after voltage was lowered from 3.0 to 2.9V. After the cell was run overnight at 3V, the current density increased to 456 mA/cm². During the whole process, ethanol reading of the cathode exhaust gas was 19.0, exceeding the analyzer's detection limit. The current density increase with time rather than decline suggests that the catalyst/cell may require a long time to reach its active working state. The results show that the EC cell does not deactivate under well-controlled conditions and instead becomes more activated with time. This may be explained by fundamental surface chemistry. Since the catalyst was stored just in ambient air, metallic surface of the copper catalyst could be oxidized during storage. Under the EC cell operation condition, the oxidized copper surface could be gradually reduced to metallic states.

TABLE 10
Stability of the cell assembly 5 loaded with cathode catalyst sheet of Cu/GDL-C

							Ethanol
Sampling		Oven T, ° C.	Cathode	Anode	Fluid T coming		in

Time,

Power

Set

gas

electrolyte

out of the cell, ° C.

Δp,

cathode

min	V	mA/cm2	point	actual	L/min	cc/min	Electrolyte	Gas	Pa	exhaust

Day 1 (start)

14:01	3.0	160	50	53.5	0.10	5.0	49.2	47.7	5,552	19.0
14:15	3.0	271	50	50.4	0.10	5.0	49.6	49.2	5,582	19.0
14:51	3.0	207	50	53.3	0.10	5.0	51.1	50.5	5,055	19.0

Day 2 (after overnight run)

8:25	3.0	425	50	50.1	0.1	5	49.1	49.2	7,119	19
9:03	3.0	243	50	52.5	0.10	5.0	51.1	50.9	6,826	19.0
9:51	3.0	294	50	52.6	0.10	5.0	51.2	50.9	6,426	19.0
10:27	3.0	220	50	52.7	0.10	5.0	51.1	50.4	5,870	19.0
12:28	3.00	329	50	53.8	0.10	5.0	50.0	49.9	7,137	19.0
12:55	3.00	326	50	51.7	0.10	5.0	50.1	51.4	7,555	19.0
13:57	3.00	385	50	50.7	0.10	5.0	51.1	51.2	6,404	19.0
16:23	2.90	290	50	49.9	0.10	5.0	50.2	50.2	6,634	19.0
16:30	2.90	283	50	50.2	0.10	5.0	50.1	50.5	6,009	19.0

Day 3 (after overnight run)

8:27	3.00	1456	50	49.9	0.10	5.0	49.5	49.4	9,633	19.0

Example V. Impact of Air Impurities in Cathode Feed Gas on the Electrochemical Cell Productivity

Cell assembly 6 was made with a new copper/carbon sheet. In this run, the oven temperature was set at 50° C., CO₂ feed gas rate was set about 0.2 liter/min, electrolyte circulation rate was set at 5.1 cc/min, and voltage was controlled at 2.5V. CO₂ gas from a beverage grade CO₂ gas cylinder was mixed with air in the 20-liter vessel to simulate air impurities in the CO₂ feed gas. Pressure gradient between the cathode and anode was maintained at 1.6 to 1.8 kPa. The cell cathode and anode outlet temperatures were controlled well around 50° C. No gas or liquid crossover occurred. When the cell performance reached steady state with a CO₂ feed gas, current density was recorded and ethanol reading in the cathode exhaust gas was measured. FIG. 14 shows impact of CO₂ purity on the current density and ethanol reading. The cell worked with CO₂ feed gas of different purity levels as evaluated. In other words, the cathode catalyst was not poisoned or deactivated by presence of air impurities in the feed CO₂ gas. However, both current density and ethanol production decreased as CO₂ purity was reduced to about 95%. Once the CO₂ purity reached about 97% level, the current density and ethanol reading was about the same as 99% beverage grade CO₂.

This example demonstrates that the cathode and anode materials used in the present cell may tolerate presence of some air impurities in the feed CO₂ gas. Thus, CO₂ captured from air may not need to be purified to a high purity level for the electrochemical conversion. However, oxygen in the cathode gas may oxidize the metallic copper catalyst and/or reaction products and its content needs to be controlled below a certain level. Oxygen content in the cathode feed gas is preferred to be less than about 2 vol %.

Example VI. Comparative Reactor Configuration—Gas/Liquid Cofeed to the Cathode

Cell assembly 7 was made with a new copper/carbon sheet in the way used in the previous example. The cell was assessed in a different reactor configuration as shown in FIG. 15. CO₂ gas and electrolyte solution was introduced to the cathode as gas/liquid two-phase flow, while the electrolyte was circulated at 5 cc/min on the anode side, same as the previous runs. The oven temperature was set at 50° C. The testing was conducted at voltage 1.5V and 2.0 V. No formation of gaseous oxygen in the anodic electrolyte was observed. No ethanol was detected at the cathode effluent with the electrolyte feed only or with both CO₂ gas/liquid electrolyte feed.

Example VII. Comparative Reactor Configuration-Flow Through Cell

Cell assembly 8 was made with a new copper/carbon sheet in the way used in the previous example. The cell was tested in a flow through reactor configuration as shown in FIG. 16. 1.0M potassium bicarbonate electrolyte solution was introduced to the cathode side at flow rate of 3 cc/min and forced to permeate through the cell stack (cathode reaction zone/anode MEA). The oven temperature was set at 50° C. The testing was conducted at 1.5 and 2.0 V. The pressure gradient between the cathode and anode was greater than 20 kPa, due to flow resistance of the porous cell stack. No gaseous production formation was observed at the anode effluent, and no ethanol was detected either. The testing result indicates that bicarbonate is not a reactant to be reduced by the present cell into CO gas or ethanol.

Example VIII. Single Cell of Larger Working Area Using Different Type of Cathode Catalyst

Essential components of this cell are same as illustrated in FIGS. 1A-2B. The inlet and outlet ports and conduits are arranged differently from what is shown in FIGS. 1A-2B. As illustrated in FIG. 17A, cathode gas inlet and outlet ports, and anode liquid inlet/outlet ports are configured on the same end plate (anodic end plate) in the example. One pair of gas flow conduits inside the cell stack are used to connect the gas flow channel inside the stack to the cathode inlet and outlet ports. Another pair of liquid flow conduits inside the cell stack are used to connect the liquid flow channels inside the stack to inlet and outlet ports. Two stainless-steel (SS) 80×80 meses are placed between the liquid flow channel and anode reaction zone, and between the gas flow channel and cathode reaction zone, respectively, to serve as an electron and pressure distributor.

The cathode and anode size end plate are made of stainless steel with the same thickness. The cathode and anode flow channels are also made of stainless steel with the same thickness and same size. The membrane separator/anodic electrode (or anodic reaction zone) is made as one body by sintering a layer of porous ceramic membrane onto a porous Nickel alloy sheet. The ceramic membrane thickness is about 0.01 mm, while the porous Nickel sheet thickness is about 0.05 mm. The ceramic/Nickel sheet size is 48 mm×140 mm. Edge areas of the sheet is covered by a PTFE film for sealing purpose to leave 35 mm×115 mm working area to contact with the cathode reaction zone. The cathode reaction zone is made by casting a catalyst slurry on the exposed area of the ceramic membrane, letting solve evaporation, and pressing a porous carbon paper of 0.31 mm thickness onto the catalyst coating. Thus, the anode reaction zone/membrane separator/cathode reaction zone are integrated into one piece. The integrated piece is assembled into a working cell as illustrated in FIG. 17A.

The catalyst slurry is made by adding 209 mg carbon black (Vulcan Xc-72, 30-60 nm) 564 mg of copper metallic power (40-60 nm), and 0.2 m1 of PTFE solution (60 wt. %) into 20 m1 of ethanol. The mixture is homogenized to form a uniform and stable suspension. 8.0 ml of the coating slurry is spread on the ceramic membrane surface uniformly. As a result, 5.64 m/cm² of a cathode catalyst layer comprising weight ratio of carbon/copper/PTFE=0.78/2.12/1.0 is deposited on the ceramic membrane surface, corresponding to volume ratio of carbon/copper/PTFE=0.76/0.50/1.0.

The cell assembly is compressed uniformly with bolts and nuts on the two end plates. The cathode and anode end plates are connected to respective negative and positive output of a DC power supply to check short circuiting. No current is detected at 0.5, 1.0, 2.0, and 2.5V. The cell passes dry quality check and is installed on the flow testing apparatus. Gas and liquid tubes are connected to respective inlet and outlet ports on the cover plate as shown in FIG. 17B. CO₂ feed gas of 0.4 liter/min is introduced into the inlet port to pressurize the cell to 3 psi while all the other ports are closed. No gas leakage has been detected. Then, liquid electrolyte solution is pumped to the anode inlet line at 50 ml/min to check any liquid leakage. The electrolyte is 1M of carbonate and bicarbonate buffer solution of pH=9.3. No gas and liquid leakage are detected. The cell is ready to conduct electrochemical reactions by providing DC current to the cell.

Different from the previous small-cell testing, the cell in this example was thermally insulated without using any oven or external heating. It was expected that Joule heating may generate enough heat to compensate heat loss to environment for the large cell.

Under constant voltage, cell productivity is assessed by current density. Variation of the current density with time is shown in Table 11 along with temperature and pressure. The gas temperature at the outlet is about 20° C. higher than the inlet. The liquid temperature at the outlet is also about 20° C. higher than the inlet. The temperature increase is attributed to Joule heating. The gas and liquid outlet temperatures are close, indicating excellent heat transfer of the cell stack. The gas side pressure is kept positive while the liquid side pressure is atmospheric. Thus, no gas and liquid crossover occurred during testing period. Upon startup, it took about 50 min for the cell to get into activated work state. The current density stabilized through overnight run.

Product formation is monitored by analyzing cathode gas effluent. The gas coming out of the cathode outlet is cooled down to room temperature to condense water vapor and liquid products. The non-condensable gas is purged by Helium gas stream to sampling line of Mass Spectra (MS) analyzer. No hydrogen formation is detected. Major products are methane and ethylene. Presence of formic acid and ethanol (m/z=46) in the non-condensable gas is minor.

To assess impacts of feed CO₂ gas humidity level on the product formation, a fraction of feed CO₂ gas is let bypass an online humidifier. FIG. 18 shows three major gaseous products detected by MS with feed CO₂ gas of different degrees of humidification. The CO₂ gas passing through the humidifier results in higher product concentration than bypassing the humidifier partially or wholly. With the present cell configuration, in principle, H₂O reactants may come out of evaporation of water from the liquid electrolyte inside the membrane pore. The present cell testing results suggest that having H₂O in the feed CO₂ gas promotes production formation rate.

TABLE 11
Performance characteristics of 40 cm2 cell using nano copper powder as cathode catalyst
(CO2 feed rate = 0.4 liter/min, electrolyte circulation rate = 50 ml/min)

	Electricity			Pg, in	Pg, out	Pl, in
	supply			(PSI)	(PSI)	(PSI)

Voltage

Current

Gas T (° C.)

Liquid T (° C.)

Gas

Gas

Liquid

Time	(V)	mA/cm2	Inlet	Outlet	Inlet	Outlet	inlet	out	inlet

11:03	2.50	61	26.9	28.8	29.2	33.1	0.40	0.07	−0.1
11:55	2.50	154	29.1	47.4	32.7	48.4	0.47	0.30	−0.1
12:23	2.50	157	29.0	48.8	29.4	47.9	0.29	0.28	−0.5
1:30	2.50	154	28.8	47.3	28.4	48.1	0.94	0.47	−0.2
2:40	2.50	143	29.8	51.8	35.1	54.6	0.45	0.33	0.2
3:20	2.50	133	29.9	54.6	34.4	55.2	0.66	0.46	−0.4
4:13	2.50	128	31.1	56.8	35.1	55.6
5:05	2.50	132	31.9	56.1	35.3	55.2	0.63	0.36	−0.7

Next day

8:05	2.50	135		47.1		46.8
9:54	2.50	138	28.6	52.9	34.8	53.7	0.41	0.33	0.1
10:58	2.50	127	28.8	53.2	34.8	53.8	0.32	0.02	0.0

Example IX. Two-Cell Stack Using Different Membranes

The same parts as used in example VII are used to assemble a two-cell stack as illustrated in FIGS. 19A and 19B. Major procedures are outlined as follows: 1) an anode end plate is placed on a flat work surface; 2) liquid flow channel plate comprising a frame and inter-connected channels is placed on the end plate; 3) an anode reaction plate comprising a gasket and porous Nickel sheet of 0.05 mm thickness is placed; 4) a membrane separator comprising a porous polymeric membrane and ceramic-coated porous Nickel sheet is placed; 5) a cathode reaction plate comprising Copper catalyst coating on a carbon paper and gasket is laid down; 6) a gas flow channel plate comprising a frame and inter-connected channels is placed; 7) a bipolar plate is placed; 8) process steps 2-6 are repeated; 9) cathode end plate is placed; and 10) the stack is compressed uniformly by bolts and nuts on the two end plates.

The inter-connected channels are flow channels connected to inlet and outlet ports on the end plate. In this example, the copper catalyst layer is painted on the polymeric membrane surface with the recipe and method as described in example VII. The catalyst layer is encapsulated with the porous carbon paper. The ceramic-coated porous Nickel sheet provides strong and durable support to the polymeric membrane. Soft polymeric membrane provides some protection function to the ceramic coating.

The two-cell assembly in this example is a stack of two identical cells in a sequence of same charging order by use of a bipolar plate, i.e., +the positive side/membrane/negative side-bipolar plate+the positive side/membrane/negative side. In FIG. 1b, the multicell are stacked in a sequence of alternate charging order without using a bipolar plate, that is, +the positive side/membrane/negative side−anode flow channel−negative side/membrane/positive side+cathode flow channel+the positive side/membrane/negative side. The basic ideas are the same-stack of an array of identical cells with gas and liquid conduits to connect the gas flow channel and liquid flow channel inside the stack to respective inlet and outlet ports on one of the end plates.

After passing dry-cell electron leakage, and gas and liquid leakage tests, the cell is installed on the flow testing apparatus to conduct electrochemical reaction tests. The cell is thermally insulated. The results are summarized in table 12. It is noted that the two cells are connected in series electrically. Thus, voltage applied to each cell is total voltage divided by 2. The current density is the same for the two cells. Gas and liquid flows for the two cells are connected in parallel. Thus, gas flow into each cell is total gas flow divided by 2. Gas outlet temperature is the same for the two cells. The same applies to liquid flow. It takes about 4 h for the cell stack to get activated and stabilized. The current density measured in this two-cell stack is less than the single cell stack, in example VII. This is likely caused by ionic transport resistance added by the polymeric membrane. The outlet gas and liquid temperatures are close each other, indicating excellent thermal conductivity of the cell stack. All the CO₂ feed gas goes through an online humidifier in this example. By maintaining the gas side at a positive pressure and leaving liquid under atmospheric pressure, no gas/liquid crossover occurred.

TABLE 12
Performance characteristics of a device
comprising two 40 cm2 cells

Electricity supply

Gas conditions

Liquid conditions

	Total		Total		Total
	voltage	Current	flow rate,	Outlet T	flow rate,	Outlet T
Time	(V)	mA/cm2	L/min	(° C.)	cc/min	(° C.)

9:46	1.00	0.0	1.0	46.6	50	47.6
10:05	2.00	0.0	1.0	46.4	50	46.8
10:15	3.00	5.5	1.0	46.7	50	44.9
10:30	3.00	3.8	1.0	45.6	50	44.7
10:37	4.00	12.5	1.0	44.2	50	46.6
10:42	4.00	8.5		43.2	50	46.7
10:57	4.00	7.8	1.0	44.7	50	47.0
11:03	4.00	8.3	1.0	45.3	50	44.9
11:23	5.00	31.3	1.0	43.7	50	44.6
12:09	5.00	36.3	1.0	46.5	50	42.8
1:02	5.00	42.3	1.0	44.9	50	41.6
2:23	5.00	58.0	1.0	46.8	50	42.3
3:35	5.00	57.8	1.0	44.4	50	41.7
3:58	5.00	57.8	1.0	44.3	50	42.4

The gaseous reaction products are analyzed using MS. The cathode and anode effluent gas is purged with Helium gas to the MS sampling point. Table 13 shows these molecules detected by MS with signal strength above the baseline. With the same anode and cathode catalyst material as well as the same electrolyte, the two-cell stack produces a lot of hydrogen gas, compared to none in the single-cell testing. Consistent with the single-cell result, CH4 and ethylene are major gaseous products. These cell testing results suggest that addition of a membrane does not only affect current density but also affects product selectivity.

TABLE 13
Gaseous molecules detected by MS in the cathode
and anode effluent by Helium gas purge

	m/z	Anode	Cathode

	H2	2	190	188
	CH4	16	12	209
	ethylene	28	11	162
	Alcohol	31	0	0
	O2	32	29	2
	Formic acid or ethanol	46	0	9

Example X Multicell Scaleup

To render rapid scaleup and low-cost production, cathode reaction zone/membrane separator/anode reaction zone may be integrated into one standardized unit-cell plate. FIG. 20A illustrates basic components of a unit cell. A supporting frame (11) provides mechanical support to the active cell and slots for process flows. Supporting grids (12) in the middle zone of the plate are used to hold the anodic membrane electrode assembly (MEA), i.e., membrane separator (4)/anode reaction zone (3), and the cathode reaction zone (5) so that the unit cell may withstand a pressure gradient between the cathode and anode sides. The supporting frame has one pair of slots for respective cathode feed introduction (13a) and cathode product collection (13b), which are arranged opposite each other on the edges of the plate. Similarly, the supporting plate provides a pair of slots for respective anode feed flow distribution (14a) and product collection (14b). The anodic MEA and cathode reaction layer are firmly fixed on the supporting plate to form one integrated body of structure. FIGS. 20B and 20C show respective cathode and anode side of the unit cell plate. On the cathode side, sealant may be used to isolate the cathode reaction zone from the environment and from the anode flow slots. On the anode side, sealant may be used to isolate the anode reaction zone from the environment and from the cathode flow slots. Shapes and sizes of the flow slots may be optimized based on specific application conditions. The working area of the unit cell may be enlarged by having an array of identical unit cells on one supporting frame. For example, if the unit cell working area is 200 mm×200 m, 6 of such identical unit cells may be configured on a support frame to obtain working area of 400 mm×600 mm.

Channel flow spacers (components 2 and 6) may be made as standard parts for stacking of the unit cell plate into a reactor working module. The channel spacer designs are illustrated in FIGS. 21A and 21Bb for respective cathode and anode flows. The cathode flow spacer (FIG. 21A) contains the following functions: supporting frame (11), flow channels in the middle portion of the plate (12), slots for cathode flows (13A, 13B), slots for anode flows (14A, 14B), and sealant (15). The flow slots and sealant match those on the cathode side of the unit cell plate (FIG. 20B). The anode flow spacer (FIG. 21B) has the structural and functional features similar to the cathode flow following functions: supporting frame (11), flow channels in the middle portion of the plate (12), slots for cathode flows (13A, 13B), slots for anode flows (14A, 14b), and sealant (15). The flow slots and sealant match those on the anode side of the unit cell plate (FIG. 20C). The spacers for the cathode and anode look similar in general features. However, those two parts are not necessarily the same, because the cathode flow is gas, and the anode flow is liquid and liquid-gas. The channel spacer plates illustrated in FIGS. 21a and 21b are equivalent to components 6 and 2 in FIG. 1, respectively.

A group of identical unit cell plates may be stacked up to form a reactor device with the flow spacers and cover plates, as illustrated in FIG. 22. The cover plate contains flow conduits to match with the flow slots on the flow spacers for connections to the fluid inlet and outlet. In the stack, each cathode flow spacer serves the two adjacent cathode reaction zones, and each anode spacer serves the two adjacent anode reaction zones. All these plates are stacked together tightly to ensure that there is no fluid leakage and there is little or no electrical contact resistance. The flow spacer in contact with the cover plate may be eliminated by using the cover plate with built-in flow channels. Cathode gas inlet and outlet ports are connected to respective introduction and discharge slots on the stack for cathode flows. Anode liquid inlet and outlet ports are connected to respective introduction and discharge slots on the stack for anode flow.

Based on the ideas and working principles of the stack design described in this invention, each component or plate may be optimized to meet the specific application needs. For example, the flow slots may be just made like a hole rather than rectangular shape. An air of holes may be located at two opposite corners of the unit cell plate and flow spacers. The two holes are connected via flow channels. A flow is introduced from one hole, uniformly distributed among the flow channels, and collected to the opposite hole.

Example XI Electrochemical Reaction Integrated with Product Separation

The process flow of a conversion plant using the electrochemical reactor of this invention is shown in FIG. 23. CO₂ gas and water are introduced to a feed conditioning vessel (C1). Humid CO₂ gas stream (2) is introduced to cathode flow channel in the EC reactor module. The cathode effluent (3) goes to a separator (SEP1) in which the desired product molecule is recovered, entrained water is recovered, and un-converted CO₂ gas along with un-desirable by products such as methane and/or hydrogen gas is recycled on a gas recycle pump (CP1). Methane and hydrogen in the recycled gas may be combusted into CO₂ and H₂O on an oxygen combustor (OC1). On the anode side, a liquid circulation pump (CP2) is used to pump the liquid electrolyte solution and remove reaction heat and byproduct oxygen gas out of the anode in the EC reactor module. The oxygen gas is separated from the liquid electrolyte along with the volatile hydrocarbon entrained by the electrolyte. The vapor-phase stream (8) is separated into oxygen gas, water, and hydrocarbon products. A heat exchanger (HX1) may be incorporated into C1 to supply the heat for liquid water vaporization, and a heat exchanger (HX-2) may be incorporated into C₂ for removal of heat from the electrolyte solution.

The EC reactor module operates at temperatures of 25 to 120° C., preferably 35 to 95° C. The pressure is controlled by about 1.1 to 2.0 bar in the cathode side and about 1.0 to 1.8 bar in the anode side. The cathode is always kept at a higher pressure than the anode. Thus, the conversion plant is a low-temperature and low-pressure system. The plant is operated with a closed loop of C and H atoms, i.e., all the water is recovered for re-usage, and all the non-product carbon molecules are recovered as CO₂.

Given the relatively low product concentrations in the cathode and anode effluents as discussed in the previous product mass transfer sections, low-cost and efficient separation of the product molecule out of the mixture may have a substantial impact on the process economics. Selective solid adsorption may be used for these separation modules. In particular, the modular adsorption unit with high productivity and energy efficiency as disclosed in MWI's prior patents may be used. The solid adsorbents have molecular selectivity so that the desired product molecule is captured as a gas mixture passes through the adsorption channels at low pressure drops. After the adsorbent is saturated, the product molecule is released by heating up the adsorbent to a higher temperature and lowering the adsorption vessel pressures. The examples of selective adsorbents for adsorption of hydrocarbon molecules are hydrophobic zeolites, MFI-type zeolite with high Si/Al ratio, carbon molecule sieves, and activated carbons.

The product gas effluent may have high water content, which may increase molecular separation energy consumption. The humid product gas stream may be dehumidified using a membrane humidity exchanger. The exchanger has an H₂O-molecular selective membrane. When the humid product gas stream flows over one side of the membrane, H₂O molecules permeate across the membrane and are swept away by a drier CO₂ gas stream. The humidified CO₂ gas may be fed to the EC cathode. In this way, water molecules are recovered without any loss of latent thermal energy.

Example XII Electrochemical Conversion of CO₂ with Direct Air Capture

The CO: conversion plant may be used to produce hydrocarbon chemicals and/or fuels using renewable electricity, such as photovoltaic (PV) panels, wind turbines, etc. CO₂ and water feeds may not be available at renewable energy locations. Transporting these feeds from point of sources to these locations could be too expensive to be economic. On other hand, presence of CO₂ and H₂O in air is ubiquitous. CO₂ and H₂O may be captured from air, using adsorption technologies. Integration of the direct air capture (DAC) of CO₂ and H₂O with electrochemical conversion enables the plant to be deployed wherever renewable energies are available without limitations by CO₂ and H₂O supplies.

FIG. 24 shows process flow diagram of an integrated plant. Adsorbents may be made as mobile modular DAC carts. CO₂ and H₂O are captured on the adsorbent by letting air flow through the adsorption channel inside the cart. After the adsorbent is saturated, the DAC cart is moved into a regeneration chamber (R1) in which the adsorbent material is heated up by a thermal fluid and residual air is removed by a rough vacuum pump. When CO₂ and H₂O desorption starts, the desorbed CO₂ and CO₂ may be pumped out of the regeneration chamber by a vacuum pump (VP1). The resulting CO₂ and H₂O are conditioned in a vessel (C1) and then introduced to the EC reactor. In the integrated plant, low grade heat rejected from the EC reactor may be upgraded using a heat pump (HP1) to supply the heat for the regeneration process. Thus, no additional heat may be needed. In the integrated unit, DAC and EC reactors may be started when renewable electricity is available. Thus, compression and storage of CO₂ is minimized.

Example. XIII Integration of the Electrochemical CO₂ Conversion Unit with Renewable Electricity

Integration of the electrochemical unit of this invention with renewable energy is illustrated in FIG. 25 with solar photovoltaic (PV) electricity as an example. When sunlight is available, electricity produced from the PV panel may be used to drive direct air capture (DAC) process to produce CO₂ and H₂O reactants and subsequent electrochemical (EC) conversion units to produce liquid oxygenated hydrocarbons. The liquid product may be transported to the customer sites for different applications. The mixture of different carbon numbers may be directly used as clean fuels, separated into individual compounds as chemical products, or upgraded into pure hydrocarbon liquid fuels. For example, the oxygenated may be dehydrated to olefins free of oxygen atoms, and the olefins may be polymerized into hydrocarbons of targeted carbon numbers.

The excess electricity may be stored in the battery stack so that the direct air capture and conversion (DACC) plant may be operated during night. Waste heat rejected by the EC stack may be upgraded by a heat pump to supply heat to DAC regeneration process. Un-converted CO₂ and un-desirable side products, such as methane and hydrogen gas, are completely recycled by oxidation reaction to produce CO₂, H₂O. Thus, the DACC plant is a closed-loop system that does not emit any greenhouse gas and does not waste any water and is a net CO₂ sinker.

The DACC plant is not limited by availability of feedstock and may be deployed wherever renewable energies are available.

Because the DACC plant operates at low temperature and low pressures, all the major pieces of equipment may be fabricated as modules (FIG. 26), similar to the PV module and lithium-ion battery stacks. A large number of identical pieces of modular equipment may be mass produced at low costs. The modules render easy transportation and installation and provide operational flexibility to change production capacity based on renewable energy outlet.

For the DAC section (FIG. 26A), solid adsorbents are made as an adsorption and heat exchanger (AHX) mobile cart. The carts are transported between the air tunnel and vacuum tunnel for respective capture and regeneration operation. The air tunnel may be assembled from a series of identical tunnel enclosures in which air is blown through adsorption channels inside the AHX cart. The air tunnel enclosure protects the AHX cart from weathering, such as snow, rain, and dusting. The vacuum tunnel may be assembled from a series of identical vacuum chamber modules. The vacuum doors are installed at the inlet and outlet of the vacuum tunnels. PV panels may be used to cover the air tunnel to efficiently utilize the land space and save the construction material.

For the EC section (FIG. 26B), the EC reactor may be built in a modular form, which is commonly used for the membrane-based separation and reaction plants in the industry. The adsorption or membrane-based product separation equipment may also be made in the modular form. A group of identical EC modules may be configured with common feed and product lines. A group of identical Sep modules may also be configurated with common feed and product lines. Both EC and Sep modules are stationary. They may be placed underneath PV panel canopy to reduce land usage and to protect the modules from weathering.

The DACC plant powered by solar PV panels does not need agricultural land, does not consume much water, does not use any fertilizers and pesticides, and does not generate environmental emissions. It is a sustainable way to produce carbon-neural hydrocarbon products and/or fuels. Based on the current electrochemical cell performance and solar PV power production efficiency, ethanol-equivalent C2+hydrocarbon fuels produced per area of the land may be 70 times of ethanol productivity from corn.

Referring to all drawings and according to various embodiments of the present disclosure, an electrochemical device for conversion of CO₂ into hydrocarbon products is provided, which comprises: a cell stack sandwiched between a first cover plate and a second cover plate, wherein the cell stack comprises at least one full cell comprising: a flow plate comprising inter-connected gas flow channels; a negatively-charged cathode catalytic reaction zone comprising a cathode; a membrane separator; a positively charged anode catalytic reaction zone comprising an anode and a porous Nickel alloy sheet of uniform pores in contact with the membrane separator; and a flow plate containing inter-connected liquid flow channels, wherein the first cover plate comprise: gas inlet ports connected with the inter-connected gas flow channels in the cell stack configured to introduce CO₂-containing feed gas into the negatively-charged cathode catalytic reaction zone; and gas outlet ports connected with the inter-connected gas flow channels in the cell stack configured to discharge of gaseous products and un-converted feed gas; wherein the second cover plate comprise: liquid inlet ports connected with the inter-connected liquid flow channels in the cell stack configured to introduce a liquid electrolyte solution into the positively charged anode catalytic reaction zone; and liquid outlet ports connected with the inter-connected liquid flow channels in the cell stack configured to discharge of oxygen product gas and liquid products; wherein the liquid electrolyte solution has pH greater than 7; wherein the electrochemical device allows operation at gas flow channel pressure higher than the liquid channel pressure at reaction temperature and under applied voltage between the cathode and the anode.

In an embodiment, the membrane separator is a porous ceramic membrane coated on the porous Nickel support sheet in the anode catalytic reaction zone at a thickness in a range from 5 to 40 μm and with pore size in a range from 2 to 100 nm. In an embodiment, the membrane separator is a thin polymeric membrane on the porous Nickel support sheet in the anode catalytic reaction zone at a thickness in a range from 0.02 to 0.2 mm. In an embodiment, the membrane separator is a polymeric membrane of thickness in a range from 0.02 to 0.2 mm on a porous ceramic coatings of a thickness in a range from 5.0 to 40 μm on the porous Nickel support sheet in the anode catalytic reaction zone. In an embodiment, the negatively-charged cathode catalytic reaction zone comprises a copper catalyst coated on a porous carbon support sheet, wherein the copper catalytic coating thickness is from about 100 to 5000 nm, and wherein the porous carbon sheet has a thickness in a range from about 0.2 to about 3.0 mm and a porosity in a range from about 40 to 90%. In an embodiment, the negatively-charged cathode catalytic reaction zone comprises a copper catalyst layer of a thickness from about 0.005 to 0.2 mm sandwiched between the membrane separator and a porous carbon support sheet, wherein the copper catalyst layer comprises copper catalyst particle sizes ranged from 2 to 200 m at volume fraction of 0.1 to 0.6, and wherein the porous carbon sheet has a thickness in a range from about 0.2 to about 3.0 mm and a porosity ranging from about 40 to 90%. In an embodiment, the negatively charged cathode catalytic reaction zone comprises a copper catalyst coated inside a porous carbon support sheet, wherein the porous carbon sheet has a thickness in a range from about 0.2 to about 3.0 mm and has a porosity in a range from about 40 to 90%. In an embodiment, the negatively-charged cathode catalytic reaction zone comprises at least one of an electron or a pressure distributor in contact with the inter-connected gas flow channels, wherein the pressure distributor is one of metal meshes, foams or screens.

In an embodiment, the positively-charged anode catalytic reaction zone comprises at least one of an electron or a pressure distributor in contact with the inter-connected liquid flow channels, wherein the pressure distributor is one of metal meshes, foams or screens. In an embodiment, the membrane separator and the positively-charged anode catalytic reaction zone are integrated into one plate of a sheet as an anodic membrane electrode assembly comprising a porous ceramic coating on a porous Nickel alloy support sheet. In an embodiment, the membrane separator, the positively-charged anode catalytic reaction zone, and the negatively-charged cathode catalytic reaction zone are integrated into a unit cell plate comprising a supporting frame that supports the unit cell plate. In an embodiment, the cell stack is stacked in sequential electrical charging order comprising a bipolar plate between two adjacent cells. In an embodiment, the cell stack is stacked in an alternative charging order comprising one of the inter-connected gas flow channels placed between the two adjacent cells. In an embodiment, the cell stack is electrically connected in series for all the cells to be charged at the same voltage differential between the cathode and the anode.

In an embodiment, the feed gas introduced from the gas inlet ports connected with the inter-connected gas flow channels in the cell stack configured to introduce CO₂-containing feed gas into the negatively-charged cathode catalytic reaction zone is uniformly distributed among all of the inter-connected gas flow channels in the cell stack. In an embodiment, the liquid electrolyte solution introduced from the liquid inlet ports connected with the inter-connected liquid flow channels in the cell stack is uniformly distributed among all of the inter-connected liquid flow channels in the cell stack. In an embodiment, the hydrocarbon products are C2+ hydrocarbons comprising at least one of ethylene, ethanol, propanol, acetone, acetone, acetate, and their mixtures. In an embodiment, the porous Nickel alloy sheet has a thickness in a range from about 40 to 200 μm, and has a 35-60% porosity, and a mean pore size of 0.2 to 2.0 μm, free of any pores above 10 μm. In an embodiment, the CO₂-containing feed gas contains water vapor of about 2 to 50 vol %. In an embodiment, the CO₂-containing feed gas contains less than 2 vol % oxygen. In an embodiment, the CO₂-containing feed gas contains is captured from air. In an embodiment, DC voltage from about 1 to about 3 V is applied to each of the at least one full cell in the cell stack. In an embodiment, the feed gas introduced from the gas inlet ports connected with the inter-connected gas flow channels in the cell stack is operated at a pressure ranging from 1.05 to 3.0 bar (absolute). In an embodiment, the liquid electrolyte solution introduced from the liquid inlet ports connected with the inter-connected liquid flow channels is operated at a pressure ranging from 1.0 to 2.0 bar (absolute).

The foregoing outlines features of several embodiments so that those skilled in the art may better understand the aspects of the present disclosure. Each embodiment described using the term “comprises” also inherently discloses that the term “comprises” may be replaced with “consists essentially of” or with the term “consists of” in some embodiments, unless expressly disclosed otherwise herein. Whenever two or more elements are listed as alternatives in a same paragraph or in different paragraphs, a Markush group including a listing of the two or more elements may also be impliedly disclosed. Whenever the auxiliary verb “can” is used in this disclosure to describe formation of an element or performance of a processing step, an embodiment in which such an element or such a processing step is not performed is also expressly contemplated, provided that the resulting apparatus or device may provide an equivalent result. As such, the auxiliary verb “can” as applied to formation of an element or performance of a processing step should also be interpreted as “may” or as “may, or may not” whenever omission of formation of such an element or such a processing step is capable of providing the same result or equivalent results, the equivalent results including somewhat superior results and somewhat inferior results. Those skilled in the art should appreciate that they may readily use the present disclosure as a basis for designing or modifying other processes and structures for carrying out the same purposes and/or achieving the same advantages of the embodiments introduced herein. Those skilled in the art should also realize that such equivalent constructions do not depart from the spirit and scope of the present disclosure, and that they may make various changes, substitutions, and alterations herein without departing from the spirit and scope of the present disclosure.