System and method for controlling carbon dioxide reactor

Abstract

To provide a system for generating liquid hydrocarbon from carbon dioxide.SOLUTION: The system comprises (a) a carbon dioxide reduction electrolysis tank comprising a membrane electrode assembly comprising an ion conductive polymer layer, and a cathode comprising a cathode catalyst layer for accelerating the chemical reduction of carbon dioxide, the cathode catalyst layer contacting with the ion conductive polymer layer, (b) a liquid hydrocarbon synthesis reactor configured to generate a mixture of liquid hydrocarbons from carbon monoxide and hydrogen, (c) a gas separation unit for removing carbon dioxide from an outlet stream to form a stream of separated carbon dioxide, and (d) a carbon dioxide recirculation loop configured to recirculate the stream of separated carbon dioxide to the cathode of the carbon dioxide reduction electrolysis tank. The system is configured to transport hydrogen and carbon monoxide from the carbon dioxide reduction electrolysis tank to the liquid hydrocarbon synthesis reactor.SELECTED DRAWING: Figure 6A

Description

Technical Field

[0001] Incorporation by Reference The PCT Claim Form is filed herewith as part of this application. Each application that claims the benefit or priority as specified in the simultaneously filed PCT Claim Form is hereby incorporated by reference in its entirety for all purposes. Description of Government Support

[0002] This invention was made with government support under contract numbers FA864920P0616 and FA8649-19-9-9026 awarded by the United States Air Force, under grant numbers DE-SC0015872, DE-SC0017725, DE-SC0018549, and DE-SC0018549 awarded by the Department of Energy, Office of Science, and under grant number 1738554 awarded by the National Science Foundation. The government has certain rights in this invention.

[0003] This disclosure generally relates to the field of carbon dioxide reactors, and more specifically, to novel and useful systems and methods for reactor control in the field of carbon dioxide reactors.

Background Art

[0004] Typical systems and methods for carbon dioxide reactor control focus on maximizing aspects related to the production of carbon monoxide (CO) and / or other carbon-containing products (CCPs), such as the ratio of CO to other reactor products (e.g., the ratio of CO to H2), CO concentration, and / or the total CO output or production rate, or adjusting them.

[0005] Accordingly, there is a need in the field of carbon dioxide reactors to create new and useful systems and methods for reactor control.

Summary of the Invention

[0006] Some aspects of the present disclosure relate to a system for producing a polycarbonate polymer. Such a system may include the following mechanisms, namely (a) a carbon dioxide reduction electrolyzer comprising a membrane electrode assembly having one or more ion-conductive polymer layers and a cathode catalyst for promoting the chemical reduction of carbon dioxide to carbon monoxide; (b) a plurality of intermediate reactors collectively configured to receive the carbon monoxide produced by the carbon dioxide reduction electrolyzer and produce one or more intermediate chemicals; (c) a polycarbonate synthesis reactor configured to receive one or more intermediate chemicals or one or more derivatives thereof and synthesize a polycarbonate polymer.

[0007] Certain aspects of the present disclosure relate to a method for producing a polycarbonate polymer. Such a method may include the following operations, namely (a) reducing carbon dioxide to carbon monoxide in a carbon dioxide reduction electrolyzer comprising a membrane electrode assembly having one or more ion-conductive polymer layers and a cathode catalyst for promoting the chemical reduction of carbon dioxide to carbon monoxide; (b) reacting the carbon monoxide produced by the carbon dioxide reduction electrolyzer in one or more of a plurality of intermediate reactors to produce one or more intermediate chemicals; and (c) synthesizing a polycarbonate polymer from one or more intermediate chemicals or one or more derivatives thereof.

[0008] Certain aspects of the present disclosure relate to a system for producing metal formates. Such a system may be characterized by the following mechanisms, namely, (a) a carbon dioxide reduction electrolyzer comprising a membrane electrode assembly having one or more ion-conductive polymer layers and a cathode catalyst for promoting the chemical reduction of carbon dioxide to carbon monoxide, (b) a formate synthesis reactor configured to receive the carbon monoxide produced by the carbon dioxide reduction electrolyzer and produce a metal formate, and (c) one or more units configured to separate and / or purify the metal formate produced by the formate synthesis reactor.

[0009] Certain aspects of the present disclosure relate to a method for producing metal formates. Such a method may be characterized by the following operations, namely, (a) reducing carbon dioxide to carbon monoxide in a carbon dioxide reduction electrolyzer comprising a membrane electrode assembly having one or more ion-conductive polymer layers and a cathode catalyst for promoting the chemical reduction of carbon dioxide to carbon monoxide, (b) reacting the carbon monoxide produced by the carbon dioxide reduction electrolyzer with a metal hydroxide to produce a metal formate, and (c) separating and / or purifying the metal formate produced in (b).

[0010] Certain aspects of the present disclosure relate to a system for producing one or more chemical compounds. Such a system may be characterized by the following mechanisms, namely, (a) a carbon dioxide capture unit configured to capture carbon dioxide from air and output carbon dioxide at a concentration higher than the concentration of carbon dioxide in the air, and (b) a carbon dioxide reduction electrolyzer comprising a membrane electrode assembly having one or more ion-conductive polymer layers and a cathode catalyst for promoting the chemical reduction of carbon dioxide to a carbon-containing reaction product. The system may be configured to provide carbon dioxide from the carbon dioxide capture unit to the carbon dioxide reduction electrolyzer.

[0011] Certain aspects of the present disclosure relate to a system for producing liquid hydrocarbons. Such a system may be characterized by the following mechanisms, namely, (a) a carbon dioxide reduction electrolyzer comprising a membrane electrode assembly having one or more ion-conductive polymer layers and a cathode catalyst for promoting the chemical reduction of carbon dioxide to carbon monoxide, and (b) a Fischer-Tropsch reactor configured to produce a liquid hydrocarbon mixture from carbon monoxide and hydrogen, and the system is configured to transport carbon monoxide and hydrogen from the carbon dioxide reduction electrolyzer to the Fischer-Tropsch reactor.

[0012] Certain aspects of the present disclosure relate to a system for producing one or more chemical compounds. Such a system may be characterized by the following mechanisms, namely, (a) a carbon dioxide reduction electrolyzer comprising a membrane electrode assembly having one or more ion-conductive polymer layers and a cathode catalyst for promoting the chemical reduction of carbon dioxide to a carbon-containing reaction product, and (b) a gas fermentation reactor configured to receive the carbon-containing reaction product produced by the carbon dioxide reduction electrolyzer and produce one or more chemical compounds.

[0013] These and other features of the present disclosure will be described in detail below with reference to the associated drawings.

Brief Description of the Drawings

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DETAILED DESCRIPTION OF THE INVENTION

[0054] The following description of the preferred embodiments is not intended to limit the present disclosure to these embodiments, but rather is intended to enable one of ordinary skill in the art to make and use the present disclosure. 1. Overview

[0055] Systems and / or methods for carbon dioxide reactor control can be configured to control aspects of reactor production, such as aspects related to the amount, concentration, and / or ratio of reactor products. An electrochemical carbon dioxide reduction cell can be integrated with any of a variety of other chemical processing systems, such as a chemical reactor, a chemical separation unit, a purification unit, etc., along with associated sensing and / or control systems. The integrated system can use an electrochemical carbon dioxide reduction cell and another chemical processing system disposed upstream, downstream, or in parallel with the electrochemical carbon dioxide reduction cell.

[0056] Examples of carbon dioxide reactants include carbon dioxide and carbon monoxide, but typically do not necessarily have to be in gaseous form. Other examples of carbon dioxide reactants include carbonate ions and compounds, bicarbonate ions and compounds, etc.

[0057] Typical systems and methods for carbon dioxide reactor control focus on maximizing aspects related to the production of carbon monoxide (CO) and / or other carbon-containing products (CCPs) (e.g., carbon-containing species (CCSs)), such as maximizing the ratio of CO to other reactor products (e.g., the ratio of CO to H2), CO concentration, and / or the total CO output amount or output rate.

[0058] However, in some applications, simply maximizing the value of a situation may not be desirable, and instead of simple maximization, any control of such a situation (e.g., dynamic or selective control of the situation to meet a value within a target situation value range) may be beneficial. For example, it may be desirable to selectively control the CO:H2 ratio of a reactor product (e.g., enabling any control within a spectrum from the highest possible CO:H2 ratio for a given system and / or a given process down to about 1:3 CO:H2 or less). Such control allows the reactor output to be used more effectively for applications such as Fischer-Tropsch processes (e.g., controlling the reactor to produce an output CO:H2 ratio of about 1:2), chemical synthesis processes, and / or gas (e.g., syngas) fermentation processes (e.g., bioreactors) (e.g., in this case the reactor output is fed directly to a subsequent input). 2. System

[0059] The system can include a carbon dioxide reactor, such as a reactor that produces carbon-containing products (e.g., CO, alkanes, alcohols, etc.) and / or hydrogen from an input containing carbon dioxide (e.g., an input stream such as a fluid stream). Exemplary carbon dioxide electrolyzers are shown in FIGS. 2A-2D. The reactor may be configured to receive a gaseous-phase carbon dioxide input and / or to perform a reaction using gaseous-phase carbon dioxide (e.g., is a gas-phase reactor), but additionally or alternatively, it can also receive a liquid-phase carbon dioxide, supercritical liquid-phase carbon dioxide, solid-phase carbon dioxide, and / or any other suitable carbon dioxide input. Although the discussion herein focuses on carbon dioxide reactors, in many cases this discussion also equally applies to carbon monoxide reactors (e.g., electrochemical carbon monoxide reduction reactors), as well as carbonate and / or bicarbonate reduction reactors. Thus, unless otherwise specified or not apparent from the context, references to carbon dioxide reactors are understood to more generally refer to carbon dioxide reactors. As indicated, the reactor may be an electrolyzer (e.g., an electrochemical reactor) such as a gas-phase polymer electrolyte membrane electrolyzer, but additionally or alternatively, it can include any other suitable reactor.

[0060] The reactor can include one or more electrodes (e.g., anodes, cathodes), a catalyst (e.g., within and / or adjacent to the cathode and / or anode), a gas diffusion layer (e.g., adjacent to the cathode and / or anode), and / or a flow field (e.g., defined within and / or adjacent to the electrode and / or gas diffusion layer, such as one or more channels defined across the gas diffusion layer and opposing the cathode). In some embodiments, the reactor includes a membrane stack or membrane electrode assembly (MEA) having one or more polymer electrolyte membranes (PEMs) that provide ion transfer between the anode and cathode of the reactor. In certain embodiments, the reactor includes a membrane stack including a cathode layer comprising a reduction catalyst and an ion-conductive polymer, a PEM membrane (e.g., a bipolar membrane, a monopolar membrane, a membrane including one or more anion conductors such as an anion exchange membrane (AEM), a proton and / or cation conductor such as a proton exchange membrane, and / or any other suitable ion-conductive polymer, a membrane including one or more buffer layers), and an anode layer comprising an oxidation catalyst and an ion-conductive polymer. The ion-conductive polymers of each layer may be the same ion-conductive polymer or different ion-conductive polymers.

[0061] In some embodiments, one or more of the catalysts (e.g., a reduction catalyst, an oxidation catalyst) can include catalyst particles such as nanoparticles (e.g., defining a porous network of the particles). One or more of the catalysts can additionally or alternatively include one or more polymer electrolytes, and optionally, the polymer electrolyte is mixed with the catalyst nanoparticles (e.g., disposed within the porous network such as being loaded into the open regions defined by the porous network). The catalyst nanoparticles can define one or more characteristic sizes (e.g., average size, median size, minimum size, maximum size, size at a particular percentile of the particle size distribution, etc.), and / or the porous network can define porosity (e.g., the ratio of the empty space within the network), density, tortuosity (e.g., characteristic path length per layer thickness, area, and / or volume such as a path through the empty space or along interconnected particles), and / or any other suitable porous network metric.

[0062] In some configurations, the bipolar MEA has the following laminated structure, i.e., a cathode layer / cathode buffer layer (anion-conducting layer) / cation-conducting layer (which may be a PEM) / anode layer. In some implementations, the bipolar MEA has a cathode layer containing an anion-conducting polymer and / or an anode layer containing a cation-conducting layer. In some implementations, the bipolar MEA has an anode buffer layer that may contain a cation-conducting material between the cation-conducting layer and the anode layer.

[0063] In some configurations, the bipolar MEA has the following laminated structure, i.e., a cathode layer / cation-conducting layer (which may be a PEM) / anion-conducting layer / anode layer. In some applications, the bipolar MEA having this arrangement is configured in a system for reducing carbonate and / or bicarbonate raw materials such as an aqueous solution of carbonate and / or bicarbonate.

[0064] In some configurations, the MEA has the following laminate structure, namely a cathode layer / anion conductive layer / anode layer. In some implementations, this MEA does not have a cation conductive layer between the cathode layer and the anode layer. In some applications, an MEA containing only an anion conductive material between the cathode and the anode is configured in a system for reducing a carbon monoxide feedstock.

[0065] In one example (Reactor Configuration A), the system includes a carbon fiber paper gas diffusion layer (e.g., Sigracet 39BC), a catalyst layer containing about 20% by weight of gold particles of about 4 nm on Vulcan carbon and an anion conductive polymer (e.g., Fumasep FAA-3), a bipolar PEM, and a flow field including a single, double, triple, or quadruple serpentine flow field or an interlocking flow field. In a specific example, the electrode defines an area of about 25 cm 2 but can define any other suitable area additionally or alternatively.

[0066] In some embodiments, the reactor includes one or more elements as described in U.S. Patent Application No. 15 / 586,182, filed on May 3, 2017, entitled "Reactor with Advanced Architecture for the Electrochemical Reaction of CO2, CO and Other Chemical Compounds", which patent is hereby incorporated by reference in its entirety. However, the reactor can additionally or alternatively include any other suitable elements in any suitable arrangement.

[0067] U.S. Patent Application Publication No. 2017 / 0321334, filed on May 3, 2017, and U.S. Provisional Patent Application No. 62 / 939,960, filed on November 25, 2019, provide additional information regarding any embodiments and / or elements of the system and / or method, which are hereby incorporated by reference in their entirety.

[0068] The carbon dioxide reduction reactor may include two or more cells or MEAs. The plurality of cells or MEAs may be arranged in a stack and may be electrically connected to each other in series and / or in parallel. Unless otherwise specified, all references in this specification to carbon dioxide reduction reactors, carbon dioxide electrolyzers, etc. embody single cell electrolyzers and multi-cell stack electrolyzers.

[0069] The carbon dioxide reduction reactor can obtain carbon dioxide from various sources. As described, examples of carbon dioxide reactants include carbon dioxide, carbon monoxide, carbonate, and / or bicarbonate. In certain embodiments, the carbonate or bicarbonate is provided in the form of an aqueous solution (e.g., an aqueous solution of potassium bicarbonate) that can be delivered to the cathode of the reduction cell. Carbonates and bicarbonates can be obtained from various sources (e.g., minerals) and / or by various reactions (e.g., reacting carbon dioxide with a hydroxide).

[0070] The system may optionally include an upstream source of carbon dioxide connected to the input of the carbon dioxide reactor of the present disclosure, which may be a biogas generation system, a corn ethanol generation system, a beer generation system, an ethanol fermentation system such as a wine generation system, a natural gas treatment system, a cement generation system, a blast furnace system capable of generating blast furnace gas, such as a steel blast furnace system, a coke oven gas generation system, a power plant system such as an oil power plant system, a natural gas power plant system, a coal power plant system, an oil refining system, an ethylene generation system, an ammonia generation system, a hydrogen generation system, such as a water gas shift system, a natural gas treatment plant (e.g., Benfield treatment), an ethylene oxide generation system, an aluminum smelting system, a liquefied natural gas (LNG) generation system, a gasifier for solid feedstocks (e.g., municipal solid waste, biomass, or coal feedstock), a reformer (e.g., steam methane reformer, autothermal reformer), a system for performing the Bouduard reaction, direct air capture (DAC) of a carbon dioxide process, the atmosphere of a planet or satellite (e.g., space atmosphere), the soil of a satellite (e.g., the soil of the Earth's moon), and / or any other system capable of generating carbon dioxide, including one or more of these. The upstream source of carbon dioxide may be directly connected to the input of the carbon dioxide reactor of the present disclosure (functioning as an input, e.g., connected to the reduction catalyst via a cathode flow field and / or gas diffusion layer, etc.), or alternatively, the upstream source may be connected to a purification system, a gas compression system, or both a purification system and a gas compression system in any order, which are then connected to the input of the carbon dioxide system of the present disclosure. Multiple purification systems and / or gas compression systems (e.g., scrubbers, etc.) may be used.

[0071] Carbon dioxide, carbon monoxide, or carbonate provided as an input to the carbon reduction reactor can have a range of concentrations depending on the structure and operating conditions of the reactor. In certain embodiments, the carbon dioxide provided to the carbon dioxide reduction reactor has a concentration of at least about 20 mol%, or at least about 40 mol%, or at least about 75 mol%, or at least about 90 mol%. In certain embodiments, the carbon dioxide provided to the carbon dioxide reduction reactor has a concentration of about 40 - 60 mol%.

[0072] The upstream source of water for the electrolytic carbon dioxide reduction reactor may be provided in various forms from any of a variety of sources such as purified tap water, purified seawater, optionally the byproduct of direct air capture of water with carbon dioxide capture, a carbon dioxide feedstock, a combustion process that can also produce fuel cell byproducts, etc.

[0073] The system may include an input portion of a downstream system capable of converting the chemical output from the carbon dioxide reactor of the present disclosure, which is connected to the output of the carbon dioxide reactor of the present disclosure. By way of example, the downstream system of the present disclosure may include a bioreactor system, a Fischer-Tropsch system, an anaerobic fermentation system, an aerobic fermentation system, a syngas fermentation system, a ketone and / or polyketone production system, a formate production system, a formate ester production system, a formamide production system, a hydroformylation system, a methanol synthesis system, an ethylene polymerization system, a phosgene production system, an isocyanate production system, a polymer (e.g., polycarbonate, polyethylene terephthalate, or polyurethane) production system, a monoethylene glycol production system, a polyethylene glycol production system, and an oxalic acid production system, and / or any other system capable of converting the chemical output from a carbon dioxide reduction reactor, including one or more of these. The carbon dioxide reactor output of the present disclosure may be directly connected to the downstream system (e.g., via a cathode flow field and / or a gas diffusion layer), and / or the carbon dioxide reactor output may be connected to a purification system, a gas compression system, or both a purification system and a gas compression system in either order, which are then optionally connected to the input of the downstream system. Multiple purification systems and / or gas compression systems may be used.

[0074] The downstream system may produce a carbon dioxide output in addition to other product outputs. The system may further include a connection between the carbon dioxide-containing output portion of the downstream system and the input portion of the carbon dioxide reactor. The carbon dioxide-containing output portion of the downstream system may be directly connected to the input portion of the carbon dioxide reactor, or alternatively, the downstream carbon dioxide-containing output portion may be connected to a purification system, a gas compression system, or both a purification system and a gas compression system in either order, which are then connected to the input portion of the carbon dioxide reactor of the present disclosure. Multiple purification systems and / or gas compression systems may be used.

[0075] The carbon dioxide reactor can produce a range of products (e.g., methane, ethylene, carbon monoxide (CO), molecular hydrogen (H2), ethanol, formate, formic acid, acetate, acetic acid, propanol, butanol, ethane, methanol) that can be used in downstream systems and processes. Different carbon dioxide reactors (e.g., including different layer stacks, catalysts and / or catalyst layers, PEM, flow fields, gas diffusion layers, cell compression configurations, and / or any other suitable aspects, etc.) can be used to obtain different reduction products (e.g., product compositions, e.g., HCR), however, additionally or alternatively, different reduction products can also be achieved by adjusting the operating parameters, and / or can be achieved in another way. Many possible downstream systems and processes emit CO2 (examples include biological utilization of methane, biological utilization of formic acid or formate, biological utilization of acetic acid or acetate, Fischer-Tropsch process, and methanol synthesis). A carbon dioxide recirculation system of a size suitable for a particular application can be used in many of these cases to return CO2 from the downstream system output to the input of the carbon dioxide reactor of the present disclosure, thereby increasing the carbon efficiency of the overall process.

[0076] The system can further include an electrical energy source connected to the carbon dioxide reactor, and the electrical energy source comprises one or more of a solar electrical energy generation system, a wind electrical energy generation system, a geothermal electrical energy generation system, a fossil fuel electrical energy generation system, or any other system capable of generating electrical energy.

[0077] A system can be used to store electrical energy in the form of chemical energy. For example, a power producer may generate excess power during off-peak usage periods. A system including a carbon dioxide reduction reactor can respond quickly to the need to consume excess power. They do not need to warm up to operate, and they can be cycled between an on state and an off state without degrading the carbon dioxide reactor. Due to the ability to quickly respond to power utilization needs, the system can function well with intermittent power sources such as solar photovoltaic energy generation systems and wind power generation systems.

[0078] One embodiment of the system can include an upstream bioreactor, a carbon dioxide reactor, and an intermittent electrical energy source. When power is available from sunlight, wind, or low off-peak demand, or other sources, a power availability detector can be used to start the carbon dioxide reactor. Further, the system can increase the output of the upstream bioreactor, for example, by raising the temperature of the upstream bioreactor and increasing the flow of nutrients to the upstream bioreactor. For other upstream carbon dioxide sources, other means may be used as needed to increase the flow of carbon dioxide to the input of the carbon dioxide reactor of the present disclosure.

[0079] Any of the systems disclosed herein may include components (e.g., sensors, systems, etc.) for measuring the state, output, and input of the system connected to the carbon dioxide reactor. Such components may include chemical property measurement systems such as gas chromatographs, mass spectrometers, infrared spectrometers, visible light spectrometers, and / or ultraviolet light spectrometers, temperature detectors, flow rate measurement sensors, power availability detectors, and / or any other monitoring system. The monitoring system can monitor parameters of the input flow and / or output flow, parameters of the components of the input flow and / or output flow (e.g., impurity concentration, carbon dioxide concentration, product concentration, etc.), and / or any other suitable parameters of the flow.

[0080] Any of the systems disclosed herein may include components for responding to conditions measured in a system connected to a carbon dioxide reactor. Such components may include a system for adjusting flow rate, temperature, power consumption, or other system parameters. The system may include one or more carbon dioxide reactors. However, the system may additionally or alternatively include any other suitable elements in any suitable arrangement. In various embodiments, one or more monitoring or sensing components are used with a control system that includes a controller having appropriately programmed or constructed logic (e.g., a processor and memory) to determine that one or more operating conditions should be changed and to cause such operating conditions to be changed. Feedforward control and / or feedback control systems may be used. 3. Method

[0081] The method may be implemented using any of the above components that include an electrochemical carbon dioxide reduction reactor, but may additionally or alternatively be implemented using any other suitable system. The method optionally includes operating the reactor under controlled process conditions (e.g., as described in more detail below) to produce a desired output (e.g., CO, H2, etc.) in a desired ratio (e.g., the molecular hydrogen to CCP ratio (HCR) and / or the CCP to molecular hydrogen ratio), and / or changing the process conditions to change the output and / or output ratio (e.g., as shown in FIG. 1).

[0082] Operating the reactor can include providing one or more inputs (e.g., gas, liquid, solid, etc.) such as carbon dioxide, carbon monoxide, a carbon dioxide source (e.g., waste gas), and / or water, reacting all or a portion of the inputs (e.g., by applying a voltage across device electrodes) to thereby produce a product, and / or removing the product from the reactor (e.g., as an output gas stream). Such reactions can include, for example, reducing carbon dioxide and / or water to produce products such as CO (and / or other CCPs, e.g., formic acid, methanol, glyoxal, methane, acetic acid, glycolaldehyde, ethylene glycol, acetaldehyde, ethanol, ethylene, hydroxyacetone, acetone, allyl alcohol, propionaldehyde, n-propanol, etc.) and / or H2. However, operating the reactor can additionally or alternatively include causing any other suitable reaction, and / or can additionally or alternatively include any other suitable element carried out in any suitable manner.

[0083] This method can include controlling the system to achieve a desired set of process conditions (e.g., scenarios), such as process conditions (e.g., scenarios) that are known to result in a desired output metric value (e.g., a desired CCP:H2 ratio, such as a CO:H2 ratio). The method can additionally or alternatively include changing process conditions (e.g., to reduce or eliminate the difference) based on the difference between the actual output and the desired output. For example, the method can include imposing an initial set of process conditions, monitoring one or more output metrics (e.g., CCP:H2 ratio), determining that the output metric is different from the target output metric (e.g., greater than or less than the target), changing one or more process conditions to reduce the output metric difference (e.g., decreasing or increasing the process condition value, such as in a condition where the output metric tends to increase or decrease as the process condition value increases), and optionally, continuously monitoring the output metric and / or continuously changing the process conditions (e.g., implementing closed-loop control of the process conditions based on the output metric).

[0084] The method can optionally include determining a target output metric that functions to determine which parameter or which scenario to target (e.g., important parameters for a given application or downstream system). One or more target output metrics can be selected for a given process. The target output metric can be an output metric associated with (e.g., predetermined therefor, indicated thereby, etc.), randomly selected, empirically determined (e.g., through iterative testing and monitoring of downstream utilization performance), optimized (e.g., based on downstream utilization operation parameters, reactor operation parameters, etc.), specified by the user, and / or determined by other means, related to the application (e.g., the applications described above such as Fischer-Tropsch).

[0085] The method can optionally include determining a target value for a target output metric, which functions to identify a target value (from a range of values). In some variations, the target value can be a maximum or minimum value (e.g., a practically achievable maximum or minimum value, a theoretical maximum or minimum value, etc.). However, the target value can additionally or alternatively not be an extreme value (e.g., it can be an intermediate value between the maximum and minimum values or a range of values between the maximum and minimum values). The target value can be associated with the application (e.g., pre-determined, pre-associated), randomly selected, empirically determined (e.g., through iterative target value selection, monitoring the performance of downstream applications, and adjusting the target value based on the performance of the application), optimized (e.g., based on the operating parameters of downstream applications, reactor operating parameters, etc.), or determined by other means. However, the target value can be any other suitable value and can be determined by any suitable method.

[0086] Under some conditions, the method may achieve more than 95% (e.g., up to 100%) carbon dioxide conversion (e.g., CO fraction yield), for example, a system operating under such conditions can achieve at least a threshold conversion metric. However, the method can additionally or alternatively include achieving carbon dioxide conversion greater than 50%, 60%, 70%, 80%, 90%, for example, carbon dioxide conversion between 10% and 100% such as 10 - 40, 30 - 50, 40 - 60, 50 - 70, 60 - 75, 70 - 85, 80 - 95, 90 - 95, 92 - 98, and / or 95 - 100%, and / or achieving any other suitable carbon dioxide conversion.

[0087] This method optionally includes providing the reactor product (or a subset thereof) to a downstream consumer of the product (e.g., as described above with respect to uses of the reactor output, as described below as described in the exemplary section, etc.). The method can optionally include modifying the reactor product after it is produced (e.g., before supplying the modified product to a downstream consumer). Modifying the reactor product can involve the product (e.g., SO from the reactor output stream x and / or NO xIt can optionally include purifying (such as removing impurities). Altering the reactor product can additionally or alternatively include, for example, mixing additional gas (and / or other substances) into the reactor output stream (and / or input stream) to achieve a desired output metric. In one variation, if the CO:H2 ratio of the reactor output differs from the desired value, the reactor output can be adjusted by mixing it with other gases (such as substantially pure CO and / or H2, another mixture of CO and H2, for example, a previously produced and stored output of the reactor, the output of a second reactor, the output of another system, and / or waste gas, etc.). For example, the CO:H2 ratio of the output stream (and / or the gas in any other part of the reactor) can be monitored (such as continuously during reactor production), and a deviation from the desired value can be compensated for by mixing with other gases (such as adding a CO-rich mixture and / or a CO-rich mixture to increase the ratio, adding H2 and / or an H2-rich mixture to decrease the ratio). This example may also include changing process conditions to correct the reactor output (as described above, for example, with respect to closed-loop control). In a second variation where an external gas source (such as the output and / or waste gas of one or more other systems like a steel mill) is supplied to a downstream consumer (such as a gas fermenter), the reactor product is used to change the CCP:H2 ratio (such as the CO:H2 ratio) of the external gas source (for example, if the CCP:H2 ratio of the external gas source differs from the desired value, mixing the reactor product to achieve the desired value). For example, based on a deviation from the desired value of the external gas source, the process conditions can be controlled to change the CO:H2 ratio of the reactor product (such as increasing the ratio in response to an external gas supply poor in CO, decreasing the ratio in response to an external gas supply rich in CO), and / or the amount of reactor product mixed into the external gas source can be controlled (for example, to achieve the desired value). However, the reactor output stream can additionally or alternatively be changed in any other suitable way or used without change.

[0088] In some examples, the method includes determining one or more metrics (e.g., operating metrics) associated with one or more upstream and / or downstream elements of the system (e.g., downstream reactors, upstream input sections, etc.). Such operating metrics can include, for example, reactor conditions such as temperature, pressure, downstream reactor and / or upstream source output metrics such as throughput, composition, purity, metrics related to other inputs to the downstream reactor such as input flow, composition, purity, reactor efficiency metrics, and / or any other suitable metrics. In such examples, the method can include changing the operation of the carbon dioxide reactor (e.g., by changing the HCR of the carbon dioxide reactor output) based on the metrics (e.g., to improve and / or maintain the operation of the downstream reactor, such as in response to changes in the upstream source, when various carbon dioxide sources are provided, etc., to improve and / or maintain a target output metric, such as HCR or reduced product concentration). However, the method can additionally or alternatively include determining any other suitable metrics and / or operating in any other suitable manner (e.g., based on the metrics). 4. Process Conditions

[0089] The process conditions can include, for example, the input carbon dioxide flow rate and / or pressure, the input gas hydration, the current density, the voltage (e.g., maintained at about 1.5 v to 3 v, additionally or alternatively, operated at less than about 1.5 v, about 2 v to 2.5 v, about 2 v to 4 v, greater than about 4 v, and / or any other suitable voltage), and / or the temperature. The process conditions can additionally or alternatively include system configurations such as the gas diffusion layer aspect, the catalyst aspect, the flow field aspect, and / or the PEM aspect. However, any other suitable process conditions can be controlled or targeted. The process conditions can be uncontrolled (e.g., as directed by an upstream system), controlled to meet a target value (e.g., the target value can be determined based on the use receiving the reactor output, the instantaneous or expected reactor operating parameters, or can be determined in other ways), or can be determined in other ways.

[0090] The process conditions can include a pressure higher than atmospheric pressure (e.g., input gas pressure, reactor pressure, etc., such as within and / or exceeding threshold pressure ranges of about 1 - 5, about 5 - 10, about 10 - 20, about 20 - 50, about 50 - 100, about 100 - 300, about 300 - 1000, about 1 - 10, about 5 - 50, about 10 - 100, about 20 - 500, and / or exceeding about 1000 atm, about 14 - 50, about 50 - 150, about 100 - 300, about 200 - 500, about 500 - 1000, about 750 - 1500, about 1000 - 3000, about 3000 - 10,000, about 10,000 - 20,000, and / or exceeding about 20,000 psi), and / or a pressure higher than that typically achievable in electrolytic cells other than gas-phase electrolytic cells. Additionally or alternatively, it can include a pressure substantially equal to 1 atm, a pressure less than about 1 atm, and / or any other suitable pressure. The process conditions can include a temperature higher than typical room temperature (e.g., within and / or exceeding threshold temperature ranges such as about 25 - 50 °C, about 40 - 60 °C, about 50 - 100 °C, about 50 - 75 °C, about 70 - 100 °C, and / or exceeding about 100 °C), and / or a temperature higher than that typically achievable in electrolytic cells other than gas-phase electrolytic cells (e.g., reactor temperature). In addition or in place of this, it can include a temperature substantially equal to room temperature (e.g., about 20 - 30 °C), a temperature less than room temperature, and / or any other suitable temperature. However, the process conditions can additionally or alternatively include any other suitable process conditions.

[0091] At high carbon dioxide flow rates, the production of CCPs such as CO can increase (e.g., due to greater availability of carbon dioxide for reduction), and thus the CCP:H2 ratio can increase (correspondingly, at low carbon dioxide flow rates, it may lead to a decrease in CCP production and the CCP:H2 ratio). In some embodiments, a higher carbon dioxide flow rate can also result in a decrease in carbon dioxide conversion efficiency, thereby diluting the output stream (e.g., syngas output) with unreacted carbon dioxide. For example, the carbon dioxide flow rate (e.g., measured at the reactor inlet) can be about 0.1 - 1000 sccm / cm2 (e.g., about 0.1 - 1, about 1 - 10, about 10 - 100, and / or about 100 - 1000 sccm / cm 2 ) can be maintained at one or more values within the range.

[0092] In a first specific example of control based on the input gas flow rate, a reactor configuration A having a triple serpentine flow field is used, the reactor pressure is maintained substantially at 120 psi, the current density is maintained substantially at 500 mA / cm 2 , and the reactor temperature is maintained substantially at 30°C. In this particular example, substantially pure carbon dioxide gas is input at various flow rates, and input flow rates of 12 sccm / cm 2 , 20 sccm / cm 2 , and 40 sccm / cm 2 (e.g., measured at the reactor inlet) result in CO:H2 ratios of approximately 1:1, 2:1.1, and 4:1, respectively.

[0093] In a second specific example of control based on the input gas flow rate, a reactor configuration A having a serpentine flow field is used, the reactor pressure is maintained substantially at 130 psi, and the current density is maintained substantially at 500 mA / cm 2 . In this particular example, when substantially pure carbon dioxide gas is input at a flow rate of 40 sccm / cm 2 , a CO:H2 ratio of approximately 8:2 is obtained, and a ratio of approximately 1:1 is obtained at a flow rate of 12 sccm / cm 2 .

[0094] Higher carbon dioxide pressure may result in an increase in CCP separation yield and / or the CCP:H2 ratio (correspondingly, when the carbon dioxide pressure is low, the CCP separation yield and / or the CCP:H2 ratio may decrease). First, an increase in carbon dioxide pressure results in greater availability of carbon dioxide for reduction, thereby increasing the total amount of CCP produced. Second, higher pressure in the catalyst can reduce the entry of water into the catalyst (e.g., from the cathode), thereby reducing the amount of water available for reduction, which can directly increase the CCP:H2 ratio and / or suppress competition for the catalytic reaction sites and / or reaction energy (e.g., thereby promoting a decrease in carbon dioxide). Thus, in some embodiments (e.g., where a high CCP partial yield and / or CCP:H2 ratio is desired), a high reactor pressure (e.g., above 100 psi, below the carbon dioxide phase transition pressure such as the critical pressure of 1070 psi) may be used. For example, the pressure of the reactor can be maintained at one or more values within the range of about 1 to 1100 psi (e.g., about 1 to 10 psi, about 10 to 100 psi, about 100 to 300 psi, about 200 to 600 psi, and / or about 500 to 1100 psi), and / or any other suitable pressure.

[0095] In a specific example of reactor pressure-based control, reactor configuration A with a single serpentine flow field is used, and substantially pure carbon dioxide gas is input at about 100 sccm / cm 2 and the current density is maintained substantially at about 150 mA / cm 2 and the reactor temperature is maintained substantially at about 20°C. In this specific example, the pressure of the reactor is maintained substantially at various pressures, and with reactor pressures of 25 psi, 50 psi, 75 psi, and 100 psi, the CO:H2 ratio becomes about 3:2, 2.4:1, 3:1, and 5:1, and the CO fraction yields become about 59%, 69%, 75%, and 84% respectively.

[0096] Increasing the hydration of the input gas can increase the depletion of water (e.g., due to a greater availability of water for reduction), and thus potentially decrease the CCP:H2 ratio. In the case of a substantially pure carbon dioxide input, only a small amount of water (which comes in almost exclusively from the cathode side of the reactor) reaches the catalyst, resulting in a higher CCP:H2 ratio. In contrast, when using a hydrated input gas, a significant amount of water from the input gas can reach and react with the catalyst. For example, the input gas hydration (e.g., the percentage of water vapor in the input gas) can be maintained at one or more values in the range of 0% (e.g., substantially pure carbon dioxide, substantially non-hydrated input gas) to 100% (e.g., 0 - 1%, 1 - 3%, 3 - 5%, 5 - 7%, 7 - 10%, 10 - 15%, 15 - 25%, 25 - 50%, 50 - 75%, and / or 75 - 100%).

[0097] In a specific example of control based on input gas hydration, reactor configuration A with a single serpentine flow field is used, the current density is maintained substantially at 50 mA / cm 2 The reactor pressure is maintained substantially at 12 psi, and the reactor temperature is maintained substantially at 20 °C. In this specific example, carbon dioxide gas with various amounts of hydration is input at 100 sccm / cm 2 A pure carbon dioxide input gas results in a CO:H2 ratio of approximately 3:2, an input gas with 12.2% hydration results in a CO:H2 ratio of approximately 1:5.67, and intermediate hydration amounts result in CO:H2 ratios between these two values.

[0098] The reactor can show CCP and H2 generation in different regimes with respect to current density. In an idealized reactor, at low current density, water reduction does not occur, all current goes towards carbon dioxide reduction, resulting in a substantially linear dependence of CO generation on current and substantially no H2 generation. On the other hand, at higher current density, additional current (e.g., above a threshold current where substantially all carbon dioxide has already been consumed) is used to reduce water, resulting in a substantially linear dependence of H2 generation on additional current and substantially constant CO generation (e.g., as shown in Figure 3A). In many typical reactors, these idealities are relaxed, but two general regimes are still shown, where CO generation increases much faster than H2 generation in the low current density regime and then approaches a plateau in the high current density regime, while H2 generation increases more rapidly (e.g., as shown in Figure 3B). This method can include controlling the generation of CO and / or H2 (e.g., controlling the CO:H2 ratio) by operating at any or all of a wide range of current densities (e.g., by controlling the operation of the reactor within the low current density regime and / or the high current density regime). In some embodiments, the use of gaseous phase input carbon dioxide can enable relatively high current densities (while reactors using aqueous carbon dioxide can be limited to lower current densities). For example, this method can be from about 1 mA / cm 2 to about 100 A / cm 2 (e.g., about 1 - 75 mA / cm 2 , about 50 - 100 mA / cm 2 , about 100 - 200 mA / cm 2 , about 200 - 500 mA / cm 2 , about 500 - 1000 mA / cm 2 , about 50 - 1000 mA / cm 2 , about 0.5 - 10 A / cm 2 , about 1 - 2 A / cm 2 , about 2 - 5 A / cm 2 , about 5 - 10 A / cm 2 , about 5 - 100 A / cm 2 , about 10 - 20 A / cm 2 2 ​, about 20 to 50 A / cm 2 , about 50 to 100 A / cm 2 , such as about 50 mA / cm 2 , about 65 mA / cm 2 , about 80 mA / cm 2 , about 90 mA / cm 2 , about 100 mA / cm 2 , about 110 mA / cm 2 , about 120 mA / cm 2 , about 130 mA / cm 2 , about 140 mA / cm 2 , about 150 mA / cm 2 , about 200 mA / cm 2 , about 300 mA / cm 2 , about 500 mA / cm 2 , about 700 mA / cm 2 , about 1000 mA / cm 2 , about 1500 mA / cm 2 including operating at a current density between such threshold values (e.g., exceeding or falling below them) and / or any other suitable current density.

[0099] In some embodiments, an increase in reactor temperature can result in a decrease in the CO:H2 ratio (e.g., due to an increase in the entry of water from the cathode, an increase in the reactivity of water, etc.). The method can include controlling the reactor temperature within an operating range such as a range between a minimum temperature (e.g., a water freezing temperature such as 0 °C) and a maximum temperature (e.g., about 40 °C, about 50 °C, about 60 °C, about 75 °C, etc., a water boiling temperature such as 100 °C) to control the CO:H2 ratio and / or any other suitable output metric.

[0100] In a specific example of control based on reactor temperature, reactor configuration A having a quadruple serpentine flow field is used, and substantially pure carbon dioxide gas is input at 70 sccm / cm 2 , and the current density is 150 mA / cm 2It is substantially maintained at, and the reactor pressure is substantially maintained at 100 psi. In this particular example, the reactor temperature is substantially maintained at various temperatures, and reactor temperatures of 26.7 °C, 35 °C, 38.7 °C, and 41.9 °C result in CO:H2 ratios of approximately 1:0.4, 2:1, 1:1.8, and 1:3, respectively.

[0101] The properties of the gas diffusion layer (GDL) can be used, additionally or alternatively, to affect CCP and / or H2 production. For example, GDL hydrophobicity can modify H2 production (e.g., by affecting water transport), a more hydrophilic GDL promotes H2 production (thereby decreasing the CCP:H2 ratio), and a more hydrophobic GDL inhibits H2 production (thereby increasing the CCP:H2 ratio). Other GDL properties such as thickness and / or pore size can also be used to modify the reactor output.

[0102] The properties of the membrane (e.g., the polymer electrolyte membrane) can be used, additionally or alternatively, to affect CCP and / or H2 production. In an example, an anion exchange membrane favorable for CCP production can be used to achieve a high CCP:H2 ratio, a cation exchange membrane favorable for H2 production can be used to achieve a low CCP:H2 ratio, and a hybrid membrane (e.g., enabling both anion and cation transport) exhibiting various anion and cation transport properties (e.g., mobility) can be used to achieve various intermediate ratios (e.g., a membrane promoting anion transport at higher ratios and cation transport at lower ratios).

[0103] The properties of the catalyst (e.g., particle size, catalyst type, etc.) can be used, additionally or alternatively, to affect CCP and / or H2 production. For example, large catalyst particles can lead to insufficient carbon dioxide transport, thereby inhibiting CCP production and potentially reducing the CCP:H2 ratio, while small catalyst particles can be advantageous for CCP production and potentially increase the ratio. The relative numbers of active sites with a high turnover frequency for hydrogen generation ("hydrogen sites") and active sites with a high turnover frequency for carbon dioxide reduction ("carbon dioxide sites") can, additionally or alternatively, depend on the catalyst particle size, with larger catalyst particles typically having a higher ratio of hydrogen sites to carbon dioxide sites and promoting H2 production, while smaller catalyst particles typically have a lower ratio and promote CO production. The catalyst type (e.g., catalyst species) can be used, additionally or alternatively, to control the reactor output, such as by using a mixture of one or more catalyst materials. A first set of catalyst materials (e.g., gold) is advantageous for carbon dioxide reduction, and a second set of catalyst materials (e.g., platinum) is advantageous for water reduction. In examples, a substantially pure gold catalyst can be used to achieve a high CCP:H2 ratio, a substantially pure platinum catalyst can be used to achieve a low CCP:H2 ratio, and various compositions of gold-platinum mixtures (e.g., alloyed particles, mixtures of gold particles and platinum particles, etc.) can be used to achieve various intermediate ratios (e.g., more gold at higher ratios, more platinum at lower ratios). The catalyst can, additionally or alternatively, include V, Cr, Mn, Fe, Co, Ni, Cu, Sn, Zr, Nb, Mo, Ru, Rh, Pd, Ag, Cd, Hf, Ta, W, Re, Ir, Hg, Al, Si, In, Ga, Tl, Pb, Bi, Sb, Te, Sm, Tb, Ce, Nd, and / or combinations thereof. The catalyst can be associated with (e.g., attached to, supported by, embedded in, adjacent to, in contact with, etc.) one or more support structures (e.g., support particles, support matrix, etc.) that can be a conductive support structure such as carbon, boron-doped diamond, and / or fluorine-doped tin oxide.However, the catalyst can additionally or alternatively include any other suitable material.

[0104] In a specific example of control based on catalyst particle size, in both cases the reactor temperature is maintained substantially at 30 °C, the reactor pressure is maintained substantially at 100 psi, there is an inter-fitting flow field, and substantially pure carbon dioxide gas is input at 10 sccm / cm 2 and the current density is maintained substantially at 500 mA / cm 2 A variant of reactor configuration A with two catalyst particle sizes is used, which is maintained substantially at these conditions. The first set of catalyst particles has a characteristic size of 4 nm (as in the case of the standard reactor configuration A), resulting in an HCR of 1:1.6 and a voltage of 3.8 V. The second set of catalyst particles has a characteristic size of 20 nm, resulting in an HCR of 1:2.8 and a voltage of 4.2 V.

[0105] Additionally or alternatively, the characteristics of reactor cell compression can be used to affect CCP and / or H2 production. In a specific example of control based on reactor cell compression, reactor configuration A is used in both cases where the reactor temperature is maintained substantially at 30 °C, the reactor pressure is maintained substantially at 100 psi, a triple serpentine flow field has substantially pure carbon dioxide gas input at 40 sccm / cm 2 and the current density is maintained substantially at 500 mA / cm 2 using two different gasket thicknesses (resulting in greater compression for the greater gasket thickness). The first gasket is 0.012 inches thick, resulting in an HCR of 1:4 and a voltage of 3.6 V. The second gasket is 0.010 inches thick, resulting in an HCR of 1:10.1 and a voltage of 3.8 V.

[0106] The characteristics of the flow field can additionally or alternatively be used to affect the CCP and / or H2 generation. In a first specific example of control based on flow field characteristics, reactor configuration A is used under two different sets of process conditions where the reactor temperature is substantially maintained at 30 °C and the reactor pressure is substantially maintained at 120 psi. In the first set of conditions, a flow field that mates with each other is used, and substantially pure carbon dioxide gas is input at 10 sccm / cm 2 and the current density is substantially maintained at 160 mA / cm 2 to obtain a CO:H2 ratio of 1.6:1. In the second set of conditions, a quadruple serpentine flow field is used, and substantially pure carbon dioxide gas is input at 40 sccm / cm 2 and the current density is substantially maintained at 120 mA / cm 2 to obtain a CO:H2 ratio of 18.5:1.

[0107] In a second specific example of control based on flow field characteristics, reactor configuration A is used under two different series of process conditions where both the reactor temperature is substantially maintained at 30 °C, the reactor pressure is substantially maintained at 100 psi, and substantially pure carbon dioxide gas is input at 40 sccm / cm 2 and the current density is substantially maintained at 500 mA / cm 2 . In the first set of conditions, a flow field that mates with each other is used, a voltage of 3.6 V is substantially maintained, and a CO:H2 ratio of 1.6:1 is obtained. In the second set of conditions, a triple serpentine flow field is used, a voltage of 3.8 V is substantially maintained, resulting in a CO:H2 ratio of 10.1:1.

[0108] However, any other suitable flow field can be additionally or alternatively used to control the reactor output, the process conditions can additionally or alternatively include any other suitable reactor conditions, and the method can additionally or alternatively include controlling the reactor output in any suitable way. 5. Impurity Resistance

[0109] In some embodiments, such as embodiments where the reactor is operated at high pressure and / or the catalyst is held at a low voltage (e.g., a negative voltage with respect to the anode), the system and / or method can achieve high resistance to impurities and / or diluted carbon dioxide inputs (e.g., compared to other carbon dioxide reactors), e.g., resistance to poisoning by impurities in the reactor input, and / or resistance to inputs diluted by species such as methane, CO, O2, and / or N2. For example, the method can include determining target process conditions (e.g., reactor configurations such as PEM type, high target reactor pressure, etc.) to achieve impurity and / or dilution input tolerances (e.g., always selecting such process conditions, selecting such process conditions according to the current and / or expected state of the reactor input section such as impurity state and / or dilution state, etc.). These impurities can include species typically present in reactor input streams (e.g., products of coal and / or natural gas combustion, e.g., output from a coal or natural gas combustion power plant) such as SO x and / or NO x and / or can include any other impurities such as ammonia, hydrogen sulfide, and mercury. In one example, the system and / or method can effectively function using an input stream containing up to 4% CO, 6% O2, 10% N2, 800 ppm NO x and / or 100 ppm SO x as well as a total of CO, O2, and N2 impurities, e.g., 10% or less.

[0110] In a specific example of dilution input tolerance, reactor configuration A with a single serpentine flow field is used, the current density is maintained substantially at 160 mA / cm 2 the reactor pressure is maintained substantially at 110 psi, the reactor temperature is maintained substantially at 20 °C, and a carbon dioxide-containing gas with various levels of dilution in methane or nitrogen is 200 sccm / cm 2It is input with. In this specific example, the reactor performance is very resistant to methane dilution up to at least 50%, and methane concentrations of 0%, 25%, and 50% result in a CO:H2 ratio between 9.5:1 and 8.5:1 and a CO fraction more than 90%. When using 75% methane, a more significant performance degradation was observed, and the CO partial yield decreased to about 84%. In this specific example, similar resistance to nitrogen dilution was observed, and nitrogen concentrations of 0%, 25%, 50%, and 75% result in a CO:H2 ratio between 9:1 and 8:1, and nitrogen concentrations up to 50% result in a CO fractional yield more than 85% (75% nitrogen concentration results in a CO fractional yield of about 81%).

[0111] In a specific example of impurity resistance, reactor configuration A with a single serpentine flow field is used, and the current density is maintained substantially at 150 mA / cm 2 , the reactor pressure is maintained substantially at 100 psi, the reactor temperature is maintained substantially at 20 °C to 25 °C, and a carbon dioxide-containing gas containing various impurities is input at 100 sccm / cm 2 It is input with. In this specific example, the reactor output metrics (e.g., CO fraction yield) under various impurity conditions are compared with the baseline reactor performance under the same conditions, except that a carbon dioxide input amount substantially free of impurities is used. In this specific example, it is shown that the reactor performance does not deviate significantly from the baseline performance for a CO concentration of 4% or less, an NO x concentration of 800 ppm or less, an SO x concentration of 120 ppm or less, or an oxygen concentration of 6% or less.

[0112] However, the system and / or method can additionally or alternatively exhibit any suitable resistance to impurities and / or dilution inputs and / or may not exhibit such resistance at all.

[0113] In certain embodiments, one or more impurities pass through the carbon dioxide reduction reactor to an output stream where they are (a) separated upstream of another chemical reactor and / or (b) passed to another chemical reactor. In embodiments where the impurities in the output stream are passed to another chemical reactor, the impurities may be used by other reactors in the chemical operations of the process. For example, hydrogen sulfide or other sulfur-containing impurities may be used by microbial species within a downstream bioreactor. 6. Selection of System Configuration

[0114] One or more system configurations may be employed based on considerations of the desired output HCR (e.g., given a particular set of process conditions and / or a range of acceptable process conditions) and / or an HCR range, etc.

[0115] In some embodiments, this includes receiving an input containing carbon dioxide in a first reactor (e.g., an electrolyzer such as a vapor-phase electrolyzer), and electrochemically generating a first reduction product (e.g., molecular hydrogen and / or one or more CCPs other than the carbon dioxide input to the first HCR) from the input (e.g., under a first set of process conditions). The selection of the first reactor design and its operating conditions includes determining a desired HCR and / or HCR range (e.g., based on downstream reactor metrics, market price metrics, efficiency metrics, and / or any other suitable metric), and selecting a system configuration (e.g., for a second reactor) based on the first HCR and / or the desired HCR such that (e.g., the second reactor optionally operates substantially under the first set of process conditions, but additionally or alternatively under any other suitable process conditions) the second reactor outputs or is capable of outputting a reduction product having an HCR closer to the desired HCR than the first HCR. For example, the configuration of the second reactor can be selected such that the second reactor generates a second reduction product from the input under conditions substantially the same as those of the first reactor (e.g., while receiving the input under the first set of process conditions), the second reduction product contains molecular hydrogen and the same CCS as the first reduction product (e.g., including substantially all species present in the first reduction product), the second reduction product defines a second HCR that is substantially different from the first HCR, and the second HCR can be selected to be closer to the desired HCR than the first HCR. In this example and / or any other embodiment described herein, the substantial difference between the first HCR and the second HCR can include a second HCR that is at least 1%, 5%, 10%, 30%, 40%, 50%, 60%, 70%, 80%, 90%, 0.5 - 5%, 2 - 10%, 5 - 25%, 20 - 50%, 40 - 80%, and / or 75 - 100% different from the first HCR (e.g., greater than or less than the first HCR) and / or different from the first HCR in some other way.

[0116] In some embodiments, selecting the system configuration can include selecting one or more aspects of the PEM to, for example, change the output HCR. Such selection can include selecting the membrane composition (e.g., different polymer species) and / or microstructure, selecting the thickness of the membrane layer, and / or selecting any other suitable aspect of the PEM. In some examples, such selection can include selecting the thickness of the anion exchange membrane and / or proton exchange membrane (e.g., a bipolar PEM with more AEM tends to produce a lower output HCR than one with more proton exchange membrane). In a first specific example, selecting a thinner AEM (e.g., thinner than the reference AEM thickness such as the thickness of the first reactor AEM, thinner than the optimized AEM thickness substantially corresponding to optimal CCP generation) can result in a reactor configured to produce a higher output HCR, and selecting a thicker AEM (e.g., thicker than the reference AEM thickness but optionally not thicker than the optimized AEM thickness) can result in a reactor configured to produce a lower output HCR.

[0117] Selecting the system configuration can additionally or alternatively include selecting one or more aspects of the reactor catalyst (e.g., reduction catalyst, oxidation catalyst), for example, to change the output HCR. In some variations, selecting the reactor catalyst aspect can include selecting the thickness of the catalyst layer (e.g., a thicker reduction catalyst tends to produce a higher HCR). In one example, selecting a thicker reduction catalyst layer (e.g., thicker than the reference reduction catalyst layer thickness such as the thickness of the first reactor reduction catalyst layer, thicker than the optimized reduction catalyst layer thickness substantially corresponding to optimal CCP generation) can result in a reactor configured to produce a higher output HCR, and selecting a thinner reduction catalyst layer (e.g., thinner than the reference reduction catalyst layer thickness but optionally less than or equal to the optimized reduction catalyst layer thickness) can result in a reactor configured to produce a lower output HCR.

[0118] Selecting the reactor catalyst aspect can additionally or alternatively include selecting the catalyst porosity (e.g., in embodiments including catalyst particles such as nanoparticles where the catalyst layer defines a porous network), for example, a more porous reduced catalyst network tends to produce a lower HCR. In one example, selecting a less porous reduced catalyst network (e.g., having a lower porosity than a reference reduced catalyst such as the porosity of the first reactor reduced catalyst network, or lower than an optimized reduced catalyst substantially corresponding to optimal CCP production) can result in a reactor configured to produce a higher output HCR, and selecting a more porous reduced catalyst (e.g., more porous than a reference reduced catalyst, but optionally not more porous than an optimized reduced catalyst) can result in a reactor configured to produce a lower output HCR.

[0119] Selecting the reactor catalyst aspect can additionally or alternatively include selecting the catalyst-to-polymer electrolyte ratio (CPR) (e.g., a higher reduced catalyst CPR tends to produce a higher HCR), for example, by selecting the degree to which the polymer electrolyte loads the porous reduced catalyst network (e.g., in embodiments where the catalyst layer includes catalyst particles (such as nanoparticles) and one or more polymer electrolytes, the catalyst particles define a porous network that includes and / or is mixed in a medium including the polymer electrolyte). In one example, selecting a higher reduced catalyst CPR (e.g., a CPR higher than a reference reduced catalyst CPR such as the CPR of the first reactor reduced catalyst network, or higher than an optimized reduced catalyst substantially corresponding to optimal CCP production) can result in a reactor configured to produce a higher output HCR, while selecting a lower CPR reduced catalyst (e.g., lower than a reference reduced catalyst, but optionally not as low as an optimized reduced catalyst) can result in a reactor configured to produce a lower output HCR.

[0120] Selecting the reactor catalyst aspect can additionally or alternatively (e.g., in embodiments where the catalyst layer includes catalyst particles such as nanoparticles) include selecting a characteristic catalyst particle size (e.g., larger particle sizes tend to produce higher HCR). In one example, selecting a larger reduced catalyst particle size (e.g., larger than the particles of a reference reduced catalyst such as a first reactor reduced catalyst, larger than an optimized reduced catalyst that substantially corresponds to optimal CCP production, etc.) can result in a reactor configured to produce a higher output HCR, and selecting a smaller reduced catalyst particle size (e.g., smaller than the particles of the reference reduced catalyst but, e.g., equal to or greater than the particles of the optimized reduced catalyst) can result in a reactor configured to produce a lower output HCR. However, the method can additionally or alternatively include selecting any other suitable reactor catalyst aspect.

[0121] This method can additionally or alternatively include selecting reactor cell compression (e.g., lower compression tends to result in higher HCR and higher compression tends to result in lower HCR), flow field, and / or any other suitable aspect of the system.

[0122] U.S. Provisional Application No. 62 / 619,996, filed on January 22, 2018, U.S. Provisional Application No. 62 / 620,109, filed on January 22, 2018, and U.S. Provisional Application No. 62 / 685,771, filed on June 15, 2018, are each incorporated herein by reference in their entirety.

[0123] The design and operating conditions of the electrolyzer can be adjusted to produce a cathode output having a particular application and particular composition. In some implementations, one or more general principles may be applied to operate in a manner that produces the required output flow configuration.

[0124] 1. Limit the availability of carbon dioxide reactant at the cathode active site and / or increase the current density at the cathode. These operating condition ranges tend to result in the following, namely (a) first, decreasing the availability of carbon dioxide reactant and / or increasing the current density increases the fraction of CO2 converted to CO (i.e., the CO:CO2 in the output stream increases), (b) at some point, further decreasing the availability of carbon dioxide reactant and / or increasing the current density makes the hydrogen ion reduction reaction more prominent (i.e., H2:CO increases). An electrolyzer that can operate with a relatively low carbon dioxide input / relatively low availability can have a flow field or gas diffusion component that limits the access of carbon dioxide to the active site of the electrolyzer cathode. In certain embodiments, a flow field design that is not mated to each other and having a long path such as a serpentine path between the source of CO2 and the cathode results in a higher CO:H2 ratio. A mated flow field forces the input gas (carbon oxide) to flow through the gas diffusion layer and then out at different locations in the flow field. A non-mated design has a long continuous path for the carbon oxide feed gas to enter and exit the cathode. The inlet side channels are spaced apart from the outlet side channels. In certain embodiments, a relatively thick gas diffusion electrode limits the transport of CO2 mass to the cathode active site and thus tends to increase the ratio of CO:CO2 and / or H2:CO.

[0125] 2. Make hydrogen ions more available at the cathode. By making hydrogen ions more available at the cathode, a cathode product stream having a relatively high H2:CO ratio can be produced. An electrolyzer configured to provide a relatively hydrogen-rich product can (a) deplete the cathode of the carbon dioxide reactant (as described in 1), (b) allow a relatively high flux of hydrogen ions to be transported from the anode, where they are generated, to the cathode, and / or (c) use a design that operates at a relatively high cell temperature. An electrolyzer that can operate with a relatively high hydrogen ion flux to the cathode can have a membrane electrode assembly (MEA) having a cation-conductive polymer and / or a mixed-ion-conductive polymer at the cathode. Alternatively or additionally, in an MEA that includes a cathode buffer layer, the layer is designed to have a relatively low hydrogen ion transport rate and / or a relatively high hydrogen ion transport rate.

[0126] 3. Reduce the availability of hydrogen ions at the cathode. By making relatively few hydrogen ions available at the cathode, a cathode product stream having a relatively high CO:H2 ratio can be produced. An electrolyzer configured to provide a relatively hydrogen-poor product can (a) provide an excess of carbon dioxide reactant to the cathode for a given current density, (b) include an MEA design that prevents hydrogen ions from reaching the cathode, and / or (c) use a design that operates at a relatively low cell temperature. High CO2 reduction product to CO2 ratio operating parameter regime

[0127] In certain embodiments, the electrolyzer is configured to produce an output stream having a CO:CO2 molar ratio of at least about 1:1 or at least about 1:2 or at least about 1:3 and actually produces it during operation. Alternatively, the high-CO output stream may be characterized as having a CO concentration of at least about 25 mol%, or at least about 33 mol%, or at least about 50 mol%.

[0128] In certain embodiments, this high carbon monoxide output concentration is obtained by operating the carbon dioxide electrolyzer in a manner that produces any one or any combination of the following operating conditions, namely, a current density of at least about 300 mA / cm2 at the cathode, a CO2 stoichiometric flow rate of about 4 or less, or about 2.5 or less, or about 1.5 or less (as described elsewhere herein), a temperature of about 80 °C or less or about 65 °C or less, a pressure range of about 75 to 400 psig, an anodic aqueous composition of about 0.1 to 50 mM bicarbonate, and an anodic aqueous pH of at least about 1.

[0129] In certain embodiments, the electrolyzer may be constructed to support a high CO:CO2 molar ratio or concentration as defined herein by using a carbon dioxide electrolyzer having any one or any combination of the following characteristics, namely, a relatively small nanoparticle cathode catalyst (e.g., having a maximum dimension of about 0.1 to 15 nm on average), gold as the cathode catalyst material, a cathode catalyst layer thickness of about 5 to 20 µm, a cathode gas diffusion layer (GDL) having a microporous layer (MPL), a cathode GDL in which PTFE is present at about 1 to 20 wt%, or about 1 to 10 wt%, or about 1 to 5 wt%, a GDL having a thickness of at least about 200 µm, a bipolar MEA having an anion exchange cathode buffer layer with a thickness of at least about 5 µm, and a cathode flow field having parallel and / or serpentine flow channels. an operating parameter regime of high reduction product (H2 + CO) to CO2 ratio

[0130] In certain embodiments, the electrolyzer is configured to produce an output stream having an (H2+CO):CO2 molar ratio of at least about 2:1 or at least about 1:2 or at least about 1:3 and actually produces it during operation.

[0131] In certain embodiments, this high reduced product output concentration is obtained by operating the carbon dioxide electrolyzer in a manner that produces any one or any combination of the following operating conditions, namely, a current density of at least about 300 mA / cm2, a CO2 stoichiometric flow rate of about 4 or less, or about 2.5 or less, or about 1.5 or less, a temperature of about 125 °C or less, a pressure of about 800 psi or less, an anodic water composition of 0 to about 500 mM bicarbonate, and an anodic water pH of about 0 to 15.

[0132] In certain embodiments, the electrolyzer may be constructed to support a high (CO + H2) / CO2 molar ratio or concentration, as defined herein, by using a carbon dioxide electrolyzer having any one or any combination of the following characteristics, namely, a nanoparticle cathode catalyst (e.g., having a maximum dimension of about 0.1 to 1000 nm on average), a transition metal as the cathode catalyst material, a cathode catalyst layer thickness of about 0.1 to 100 µm, a cathode gas diffusion layer, with or without a microporous layer (MPL), a GDL having about 0 to 70 wt% PTFE, a GDL having a thickness of about 10 to 1000 µm, a bipolar MEA having an anion exchange cathode buffer layer with a thickness of about 0 to 100 µm. Hydrogen-rich product gas flow operating parameter regime

[0133] In certain embodiments, the carbon dioxide electrolyzer is configured to produce an output stream having an H2:CO molar ratio of at least about 1:1 and actually produces it during operation.

[0134] In certain embodiments, such a hydrogen-rich output concentration is obtained by operating the carbon dioxide electrolyzer in a manner that produces any one or any combination of the following operating conditions, namely, a current density of at least about 300 mA / cm2, a CO2 mass transfer stoichiometric flow rate to the cathode of at most about 2, a temperature of at least about 65 °C or at least about 80 °C, a pressure range of about 75 to 500 psig, a pure water or an anode water composition of at least about 50 mM bicarbonate, and an anode water pH of about 1 or less.

[0135] In certain embodiments, the electrolyzer may be constructed to support a hydrogen-rich molar ratio or concentration as defined herein by using a carbon dioxide electrolyzer having any one or any combination of the following characteristics, namely, a relatively large nanoparticle cathode catalyst (e.g., having an average maximum dimension of at least about 80 nm), silver, palladium, or zinc as the cathode catalyst material, a cathode catalyst layer thickness of about 5 μm or less or a thickness of about 25 μm or more, a cathode gas diffusion layer without a microporous layer (MPL), a cathode GDL in which PTFE is absent or contains at least about 20 wt% PTFE, a cathode GDL having a thickness of about 200 μm or less or about 500 μm or more, a bipolar MEA having an anion exchange cathode buffer layer with a thickness of about 0 - 5 μm. The operating parameter regime of high reduction products to hydrogen production stream

[0136] In certain embodiments, the carbon dioxide electrolyzer is configured to produce an output stream having a CO:H2 molar ratio of at least about 2:1 and actually produces it during operation.

[0137] In certain embodiments, such product-rich output concentrations are obtained by operating the carbon dioxide electrolyzer in a manner that produces any one or any combination of the following operating conditions, namely, a current density at the cathode of at least about 300 mA / cm2, a stoichiometric flow rate of CO2 to the cathode of at least about 1.5, or at least about 2.5, or at least about 4, a temperature of about 80 °C or less, a pressure in the range of about 75 to 400 psig, an anodic aqueous composition of bicarbonate from about 0.1 mM to 50 mM, and an anodic aqueous pH greater than about 1.

[0138] In certain embodiments, the electrolyzer may be constructed to support a product-rich molar ratio or concentration as defined herein by using a carbon dioxide electrolyzer having any one or any combination of the following characteristics, namely, relatively small nanoparticle catalysts (e.g., having a maximum dimension of about 0.1 - 15 nm on average), gold as the cathode catalyst material, a cathode catalyst layer thickness of about 5 - 20 µm, a cathode gas diffusion layer having a microporous layer (MPL), a cathode GDL in which PTFE is present at about 1 - 20 wt%, or about 1 - 10 wt%, or about 1 - 5 wt%, a cathode GDL having a thickness of at least about 200 µm, and a bipolar MEA having an anion exchange layer with a thickness of at least about 5 µm. Stoichiometric flow rate

[0139] Considering that the molar flow rate can be at least partially determined by the current supplied to the cell, the molar flow rate may be associated with the current. As an example, the molar flow rate of carbon dioxide in the input stream can be defined in terms of the flow rate per reaction unit expected for a given current. As used herein, the term "stoichiometric" flow rate refers to a fraction or multiple of the flow rate of the reactant carbon dioxide necessary to fully utilize all of the current at the cathode, assuming that the reduction reaction of carbon dioxide is 100% efficient for a given reaction at the cathode. A flow rate of carbon dioxide having a stoichiometric value of "1" is the flow rate necessary to consume all of the electrons given at the cathode and is less than that in a given reduction reaction at the cathode. In other words, the stoichiometric flow rate is the amount of excess (or deficit) reactant present in excess of (or less than) the amount that could theoretically react if the current efficiency of a given reaction were 100%.

[0140] In the carbon dioxide reduction reaction (CO2 + 2H + + 2e- → CO + H2O) that produces carbon monoxide in an acidic environment, a carbon dioxide flow rate with a stoichiometric value of 1 provides 1 mole of carbon dioxide per 2 moles of electrons provided by the cell. In other words, a cell having a current that provides 2 moles / second of electrons and a carbon dioxide flow rate that provides 1 mole / second of carbon dioxide molecules has a stoichiometric flow rate of 1. At the same current and a flow rate of 0.5 moles of carbon dioxide per second, the cell has a stoichiometric flow rate of 0.5. Also, at the same current but a flow rate of 1.5 moles of carbon dioxide per second, the cell has a stoichiometric flow rate of 1.5. The molar flow rate required to achieve a stoichiometric flow rate of 1 can be calculated, Stoichiometric flow rate (sccm) = [60 (s / min) * Molar gas volume at STP (mL / mol)] / [Faraday constant (C / mol-) * #e-'s / mol CO2] * Amperes of current supplied to the electrolyzer. The total amperes of current can be calculated from the current density, the area of the electrolyzer cell, and the number of cells within the electrolyzer, Amperes of current = Current density × Area of the electrolyzer cell × Number of cells.

[0141] In one example, an electrolyzer with a current density of 500 mA / cm for the electrochemical reduction of CO2 to CO 2 has a total current of 50 A and a reaction area of 100 cm 2 . Since the reaction requires 2 moles of e- / mole of CO produced, the stoichiometric flow rate 1 is [60 * 22,413] / [9,6485 * 2] * 50 = 348.4 sccm. In this example, the stoichiometric flow rate of 0.5 is as follows. 0.5 * 348.4 = 174.2 sccm Also, the stoichiometric flow rate 2 is 2 * 348.4 = 696.8 sccm.

[0142] In another example of a cell for producing ethylene from carbon dioxide, 12 moles of electrons are required to reduce 2 moles of carbon dioxide to 1 mole of ethylene. The stoichiometric flow rate of a 3-cell 1500 cm2 electrolyzer with a current density of 300 mA / cm2 is [60 * 22,413] / [96,485 * 6] * 1350 = 3,136 sccm.

[0143] The following examples were carried out to illustrate the influence of specific electrolyzer design and operating parameters on the molar ratio of gases in the cathode output stream.

[0144] All examples used 20 wt% Au / Vulcan XC-72R containing 4 nm Au particles in the cathode catalyst layer. All examples used 100% CO2 input without humidification as the input to the cathode of the electrolyzer. All examples used a bipolar MEA with an anion exchange polymer electrolyte adjacent to the cathode layer. The anion exchange polymer had a backbone repeating unit consisting of three aryl groups and a methylene carbon with a CF3 pendant group and an alkyl quaternary ammonium pendant group. The polymer used bicarbonate, carbonate, hydroxide, and / or bromide as the counterion to the quaternary ammonium group (Orion Polymer And Membrane, Cohoes, New York). If all IrOx catalysts are used at the anode for water oxidation, IrRuOx can also be used. The flow rate of water to the anode ranges from 4 L / min to 40 mL / min. All are single cells

[0145] Example 1 Composition of the output stream: 30% CO, 20% H2, 50% CO2 (3:2 CO:H2 ratio) Current density: 300 mA / cm 2 CO2 input flow rate: 400 sccm Cell temperature: 50 °C Cell area: 100 cm2 Type of flow field: Interdigitated type GDL type: Sigracet29BC Au metal loading: 0.3 mg / cm2 Catalyst layer thickness: 15 um AEM layer thickness: 12 um Type and thickness of the membrane: Nafion117, thickness 183 um

[0146] Applications 1. It can be directly supplied to a gas fermentation reactor, or CO2 can be removed to increase the CO + H2 concentration, or H2 can be removed or added to change the CO:H2 ratio, affecting the product or combinations thereof. 2. H2 may be added to produce F - T feedstock, and it is not necessary to remove CO2 for all reactor designs, but in some reactors, CO2 can be removed to obtain higher activity.

[0147] Example 2 Composition of the output stream: 50% CO, 5% H2, 45% CO2 (10:1 CO:H2 ratio) Current density: 400 mA / cm 2 CO2 input flow rate: 110 sccm Cell temperature: 45 °C Cell area: 25 cm2 Type of flow field: Serpentine type GDL type: Sigracet39BC Au metal loading: 0.27 mg / cm2 Catalyst layer thickness: 14 um AEM layer thickness: 14 um Membrane type and thickness: Nafion115, thickness 127 um

[0148] Applications 1. It can be directly supplied to a gas fermentation reactor, or remove CO2 to increase the CO + H2 concentration, or remove or add H2 to change the CO:H2 ratio, affecting the product or a combination thereof. 2. H2 may be added to produce the raw materials for F-T. It is not necessary to remove CO2 for all reactor designs, but in some reactors, removing CO2 can obtain higher activity. 3. In some cases, it may be necessary to remove CO2 instead of H2 for formate production. 4. For polycarbonate production, it may be necessary to remove CO2 + H2 to a total concentration of less than about 2%.

[0149] Example 3 Composition of the output stream: 20% CO, 20% H2, 60% CO2 (1:1 CO:H2 ratio) Current density: 300 mA / cm 2 CO2 input flow rate: 400 sccm Cell temperature: 50 °C Cell area: 100 cm2 Type of flow field: Interdigitated GDL type: Sigracet29BC Au metal loading: 0.3 mg / cm2 Catalyst layer thickness: 15 um AEM layer thickness: 12 um Membrane type and thickness: Nafion117, thickness 183 um

[0150] Applications 1. H2 may be added to produce the raw materials for F-T. It is not necessary to remove CO2 for all reactor designs, but in some reactors, removing CO2 can obtain higher activity.

[0151] Example 4 Composition of the output stream: 35% CO, 35% H2, 30% CO2 (CO:H2 ratio of 1:1) Current density: 300 mA / cm 2 CO2 input flow rate: 60 sccm Cell temperature: 50 °C Cell area: 25 cm2 Type of flow field: Interdigitated type GDL type: Sigracet 29BC Au metal loading: 0.32 mg / cm2 Catalyst layer thickness: 15.6 µm AEM layer thickness: 14 µm Type and thickness of the membrane: Nafion 115, thickness 127 µm

[0152] Applications 1. It can be directly supplied to a gas fermentation reactor, or CO2 can be removed to increase the CO + H2 concentration, or H2 can be removed or added to change the CO:H2 ratio, affecting the product or combinations thereof. 2. H2 may be added to produce F - T feedstock. It is not necessary to remove CO2 for all reactor designs, but in some reactors, removing CO2 can obtain higher activity.

[0153] Example 5 Composition of the output stream: 55% CO, 10% H2, 35% CO2 (CO:H2 ratio of 5.5:1) Current density: 300 mA / cm 2 CO2 input flow rate: 60 sccm Cell temperature: 50 °C Cell area: 25 cm2 Type of flow field: Serpentine type GDL type: Sigracet 39BC Au metal loading: 0.24 - 0.3 mg / cm2 Catalyst layer thickness: 13 - 15 µm Thickness of the AEM layer: 12 - 14 µm Type and thickness of the membrane: Nafion 117, thickness 183 µm

[0154] Applications 1. It may be directly supplied to the gas fermentation reactor, or CO₂ may be removed to increase the CO + H₂ concentration, or H₂ may be removed or added to change the CO:H₂ ratio, which may affect the product or a combination thereof. 2. H₂ may be added to produce the raw material for F-T. It is not necessary to remove CO₂ for all reactor designs, but in some reactors, CO₂ can be removed to obtain higher activity. 3. In the case of formate production, it is necessary to remove CO₂, but H₂ is not removed. 4. For polycarbonate production, it is necessary to remove CO₂ + H₂ to a total concentration of less than about 2%.

[0155] Example 6 Composition of the output stream: 15% CO, 1% H₂, 84% CO₂ (15:1 CO:H₂ ratio) Current density: 600 mA / cm 2 CO₂ input flow rate: 500 sccm Cell temperature: 50 °C Cell area: 25 cm² Type of flow field: Serpentine type GDL type: Sigracet 39BC Au metal loading: 0.24 - 0.3 mg / cm² Catalyst layer thickness: 13 - 15 µm Thickness of the AEM layer: 12 - 14 µm Type and thickness of the membrane: Nafion 117, thickness 183 µm

[0156] Applications 1. For formate production, it may be necessary to remove CO₂ rather than H₂. 2. For polycarbonate production, it may be necessary to remove CO₂ + H₂ to a total concentration of less than about 2%.

[0157] In various embodiments, the oxygen produced at the anode of the carbon dioxide electrolyzer is used in an integrated process. By way of example, the oxygen produced in the electrolyzer may be used in a partial oxidation gasification process, an aerobic fermentation process, an electrolysis process using an oxygen depolarized electrode, and the like. In one example of an integrated scheme, a system having a Fischer-Tropsch reactor produces synthesis gas as an input to the Fischer-Tropsch reactor and oxygen as an input to a reactor for gasification of biomass, which can use a carbon dioxide electrolyzer configured to also produce synthesis gas for input to the Fischer-Tropsch reactor. 7. Integrated scheme

[0158] Additional information regarding optional embodiments and / or elements of the system and / or method is provided below.

[0159] The product gas from the carbon dioxide reactor of the present disclosure can be used in one or more downstream processes. For example, the carbon dioxide reactor of the present disclosure configured for synthesis gas production can output a stream of CO, H2, and / or CO2. Aerobic and anaerobic fermentation

[0160] This output stream can be fed to the input section of the bioreactor, where microorganisms (e.g., Clostridium Autoethanogenum, Clostridium Carboxidovorans, Clostridium Ijungdahlii, Clostridium Ragsdalei, Clostridium thermoaceticum, Clostridium thermoautotrophicum, Eubacterium Limosum, Peptostreptococcus productus, Butyribacterium methylotrophicum, acetogens, E. coli, etc.) use the energy of some of the carbon contained in CO, H2, and / or CO and CO2 to produce one or more bioproducts (e.g., ethanol, acetic acid, butanol, butyric acid, methane, etc.). The unused carbon can be released from the output section of the downstream bioreactor (e.g., as CO2, optionally together with water vapor and / or other volatile compounds).

[0161] The CO2 released as the output of the downstream bioreactor can optionally be recycled back to the input of the CO2 reactor of the present disclosure (e.g., to increase the carbon efficiency of bioproduct production, to control the operation of the carbon dioxide reactor, etc.). In some embodiments, it may be desirable to treat this CO2 before it enters (e.g., re-enters) the CO2 reactor of the present disclosure. For example, water vapor may be removed, volatile products that inhibit the function of the carbon dioxide reactor may be removed, and / or the CO2 may be pressurized to a level desirable for the operation of the carbon dioxide reactor of the present disclosure. The carbon dioxide exiting the bioreactor may be near atmospheric pressure and / or may have any other suitable pressure, and typical carbon dioxide reactor pressures may be in the range of 20 psi to 800 psi, 50 psi to 400 psi, 100 psi to 500 psi, and / or any other suitable range. In some examples, the water vapor is removed by a phase separator and / or a desiccant (e.g., a phase separator followed by a desiccant). In some examples, the volatile products are removed by oxidation, adsorption onto a suitable adsorbent, and / or condensation. A CO2 compressor can be used to raise the pressure of the CO2 to a pressure suitable for the carbon dioxide reactor. If the carbon dioxide reactor can operate with low-pressure CO2 and is not inhibited by water vapor or volatile compounds found in the CO2 stream output from the downstream bioreactor, the system can then be simplified and unnecessary purification and compression systems and processes can be removed.

[0162] For each liter of the culture medium in the downstream bioreactor, a flow rate in the range of about 1 sccm to 1000 sccm or about 1 sccm to 2000 sccm or about 10 sccm to 500 sccm, or any other suitable range of the gas from the output of the carbon dioxide reactor may be desirable. For each liter of the culture medium in the downstream bioreactor, the CO2 released can be in the range of about 1 sccm to 2000 sccm or about 10 sccm to 1000 sccm or about 10 sccm to 500 sccm, or any other suitable range. For each liter of the culture medium in the downstream bioreactor, the water vapor in the output gas stream exiting the bioreactor can be about 1 to 2 volume % of the stream, about 2 to 5 volume % of the stream, about 5 to 10 volume % of the stream, about 10 to 25 volume % of the stream, about 25 to 50 volume % of the stream, about 50 to 90 volume % of the stream, and / or any other suitable amount. The volatile products exiting the downstream bioreactor may constitute less than about 0.1%, less than about 0.5 volume %, less than about 1 volume %, less than about 4 volume % of the stream, and / or any other suitable amount.

[0163] Some microbial processes can use the syngas produced by the carbon dioxide reactors of the present disclosure. The syngas output stream of CO, H2, and optionally CO2 can be used as a feedstock for a downstream bioreactor in which a microbial process is carried out to produce a range of useful compounds such as ethanol, acetic acid, butanol, butyric acid, acetone, methane. The syngas stream itself does not necessarily contain all the nutrients required for the growth of the microorganisms within the downstream bioreactor. The addition of other nutrients to the bioreactor may be necessary for the microorganisms to grow and produce products. Examples of suitable microorganisms include Clostridium autoethanogenum, Clostridium carboxidovorans, Clostridium ljungdahlii, Clostridium ragsdalei, Clostridium thermoaceticum, Clostridium thermoautotrophicum, Eubacterium limosum, Peptostreptococcus productus, Butyribacterium methylotrophicum, acetogens, and / or E. coli.

[0164] One nutrient that can be particularly difficult to introduce into the downstream bioreactor is sulfur. Many microorganisms require sulfur for certain amino acid synthesis and enzymatic processes. The carbon dioxide reactors of the present disclosure that are resistant to sulfur can simplify the addition of sulfur to the downstream bioreactor (e.g., in addition to providing syngas to the downstream bioreactor). H2S, SO2, and / or other sulfur oxides (SO xSulfur in the form of one or more sulfur-containing species (SCS), such as, may be present in the CO2 gas supplied to the input of the carbon dioxide reactor of the present disclosure. H2S can pass through the carbon dioxide reactor of the present disclosure unchanged and exit with the syngas output stream. SCS (e.g., SO2 and / or SO x ) may pass through unchanged and / or may be converted to one or more other SCS (e.g., H2S) and may be output with the syngas output stream. Then, syngas further containing sulfur species (e.g., H2S, SO2, and / or SO x ) can be supplied to the input of a downstream bioreactor (e.g., without the need for additional sulfur nutrients). The sulfur species concentration can be in the range of about 1 ppm to 10 ppm, about 5 ppm to 50 ppm, about 5 ppm to 100 ppm, about 10 ppm to 200 ppm, about 20 ppm to 1000 ppm, and / or any other suitable range.

[0165] In some embodiments, the carbon dioxide reactor can be coupled to one or more gas fermentation reactors (e.g., downstream of the carbon dioxide reactor, such as receiving one or more products of the carbon dioxide reactor). The method can optionally include controlling the operation of the reactor based on this coupling to optimize carbon efficiency and / or energy efficiency. As seen in the energy balance shown in Table 1, with pure CO as the input, acetate is the most energy-efficient and often shows the highest selectivity for the desired final product. However, for several reasons, an integrated electrochemical gas fermentation system may be designed to utilize hydrogen-containing syngas.

Table 1

[0166] Using CO for most or all of the electron transfer chemical reactions within a downstream bioreactor will typically result in the production of CO2, which can then be released into the output stream of the bioreactor. Typically, as the ratio of hydrogen in the syngas increases, the amount of CO2 produced decreases, and the CO2 byproduct can be stoichiometrically removed beyond a certain ratio of hydrogen to carbon monoxide. For example, in the case of gas fermentation to ethanol, a CO:H2 ratio of less than about 1:2 will typically result in the incorporation of all input carbon into the ethanol end product. Thus, by adjusting the CO:H2 ratio in the output stream of the carbon dioxide reactor of the present disclosure, the operator can optimize carbon efficiency (e.g., to minimize CO2 emissions) by shifting towards more H2 production and / or optimize energy efficiency by shifting towards higher CO production. Monitoring input costs such as electricity prices or incentives for carbon utilization can inform the optimal operating parameters at any given time. Adjusting production in this way can also change the output, for example, by driving towards more ethanol production (e.g., higher CO) or more acetate production (e.g., higher H2). Monitoring the market price of the output can inform the optimal operating parameters at any given point in time (e.g., the operating parameters are determined based on the market price to optimize the market price of the product or to optimize the total profit from reactor operation).

[0167] However, the system and / or method can additionally or alternatively include any other suitable elements.

[0168] Using a system consisting of a CO2 electrolyzer and an aerobic fermentation reactor, products such as proteins, polyhydroxyalkanoates, acetone, isopropanol, ethanol, and other products can be produced. The CO2 electrolyzer takes in CO2, water, and electricity and outputs a stream of oxygen and a separate stream containing one or more carbon-based products derived from CO2, hydrogen, unreacted CO2, and water. The carbon-based products can be CO, methane, ethylene, or a mixture of these compounds with other carbonaceous compounds.

[0169] The oxygen and carbon product streams can be fed to an aerobic gas fermentation system containing yeast, Escherichia coli, or other microorganisms that can metabolize the compounds in the gas stream to produce the desired products. The output of the aerobic gas fermentation reactor typically includes a stream containing at least one bioproduct in the liquid phase and a gas stream containing CO2, water vapor, non-metabolized components of the gas feedstock, and other volatile compounds such as trace hydrocarbons or H2S produced during fermentation. The CO2 in this gas stream can be recycled back to the inlet of the CO2 electrolyzer in the same way as in an anaerobic fermentation process, but O2 removal may be necessary if the O2 concentration in the stream is higher than about 5%, 1%, or 0.25%.

[0170] Various microbial metabolic pathways may be utilized in a gas fermentation reactor configured to receive the products of an electrolyzer for carbon dioxide. The design and operation of the electrolyzer is consistent with the metabolic pathway and thus the required inputs of the pathway. One exemplary pathway is the Wood-Ljungdahl pathway (WLP), which has a set of biochemical reactions that can utilize CO, CO2, formate, methanol, H2, and / or other single-carbon compounds to produce acetyl coenzyme A (acetyl-CoA). Acetyl-CoA is a molecule that is used as a carbon and energy source for microorganisms. The Wood-Ljungdahl pathway is a natural metabolic pathway found in acetogenic microorganisms such as Clostridium ljungdahlii, but other microorganisms such as Escherichia coli can be genetically engineered to have this pathway. Acetyl-CoA can be used by microorganisms to build cell mass and / or can also be used as a starting molecule for other biochemical pathways to produce other bioproducts such as acetone and ethanol. The biochemical pathways that utilize acetyl-CoA to produce products can be inherent to the organism or may be added by genetic engineering so that the microorganism produces the desired product.

[0171] Microorganisms may utilize multiple metabolic pathways simultaneously. For example, sugars (e.g., glucose) can be metabolized by microorganisms via glycolysis or non-oxidative glycolysis pathways simultaneously as an input to the WLP. Biomolecules (e.g., ATP / ADP, NADH / NAD+) can be produced in one metabolic pathway and used in another.

[0172] WLP does not produce ATP. When cells are ATP-deficient, they promote acetate production from acetyl-CoA. To produce ATP, cells can directly utilize H2 via a flavin-based electron bifurcation pathway to generate a proton concentration gradient across their cell membranes. This gradient can promote ATP production via ATP synthase. When cells have available ATP, acetyl-CoA can be converted into various desired products via either natural or genetically engineered biochemical pathways. The ratio of H2:CO in the gas stream can determine the availability of ATP and can determine the carbon efficiency of the bioprocess.

[0173] In some embodiments, an aerobic fermentation process using methane and carbon dioxide inputs produces 3-hydroxypropanoic acid. In some exemplary systems, a carbon dioxide electrolyzer is configured to supply methane and carbon dioxide to a gas fermentation reactor that produces, among other products, 3-hydroxypropanoic acid.

[0174] In another example, a carbon dioxide electrolyzer is configured to produce outputs of carbon monoxide, carbon dioxide, and hydrogen, which may be processed to adjust the ratio of these gases before being delivered to a gas fermentation reactor that further receives sugar (e.g., glucose) as an input and reacts these components to produce acetone.

[0175] Figure 4 shows a gas fermentation system 401 comprising a carbon dioxide reduction electrolyzer 403 upstream of a gas fermentation bioreactor 405 and a recirculation unit 413 configured to recirculate the output of the bioreactor 405 to the input of the electrolyzer 403.

[0176] As shown, electrolyzer 403 is configured to receive water and carbon oxide (carbon dioxide in this example) as reactants and electricity, and to drive an anodic reaction and a cathodic reaction. The input to electrolyzer 403 is provided at a relatively high pressure (at least exceeding atmospheric pressure). The anodic reaction of the electrolyzer produces oxygen, which is used as an input to bioreactor 405 only when aerobic fermentation is used. The output from the cathodic side of electrolyzer 403 includes unreacted carbon dioxide along with one or more carbon-containing products such as hydrogen gas, carbon monoxide, and / or methane. These outputs are provided at a relatively high pressure.

[0177] System 401 is configured to transport the cathodic side output of electrolyzer 403, optionally along with anodically generated oxygen, to bioreactor 405. System 401 is configured to reduce the pressure of some or all of the electrolyzer products before or during delivery to bioreactor 405. In certain embodiments, system 401 includes a gas purification unit (not shown) between electrolyzer 403 and bioreactor 405. The gas purification unit is configured to adjust the concentration of the electrolysis products before the electrolyzer products enter bioreactor 405. In some cases, the purification unit reduces the concentration of carbon dioxide in the gas stream. In some cases, the purification unit increases the concentration of carbon monoxide in the gas stream.

[0178] Bioreactor 405 is configured to convert the input, optionally along with other inputs such as sugars, into the desired bioproducts and by-products. In the example shown, the by-products are relatively volatile or separable from the desired products and can thus be sent to recycle unit 413. In the illustrated embodiment, system 401 is configured to convey carbon dioxide, water vapor, and optionally other volatile compounds from the bioreactor to recycle unit 413 at low pressure.

[0179] The recirculation unit 413 comprises one or more separation units or removal units and a carbon dioxide compressor. In the illustrated embodiment, the recirculation unit 413 is configured to first receive the bioreactor vapor outputs and send them to water and optionally the oxygen separation unit 407. The output of the separation unit 407 may be a mixture of low-pressure carbon dioxide and optionally volatile compounds. The recirculation unit 413 is further configured to send carbon dioxide and volatile substances to the volatile compound removal unit 409, which removes the volatile substances and outputs low-pressure carbon dioxide. The recirculation unit 413 also includes a carbon dioxide compressor 411 configured to receive the low-pressure carbon dioxide from the removal unit 409 and pressurize it to a level suitable for input to the electrolyzer 403.

[0180] In certain embodiments, the recirculation unit 413 includes one or more carbon dioxide capture units that include an adsorbent for capturing carbon dioxide during a first stage and releasing carbon dioxide during a second stage. The separation unit 409 may include or be configured to cooperate with such carbon dioxide capture units. Examples of such capture units are provided in the description of the direct air capture units described herein.

[0181] In the illustrated embodiment, the system 401 is configured to transport the pressurized carbon dioxide from the recirculation unit 413 to a location upstream of the cathode-side inlet to the electrolyzer 403, where the carbon dioxide is mixed with the pressurized feed carbon dioxide.

[0182] In certain embodiments, depending on the bioreaction being performed, the carbon dioxide electrolyzer located upstream from the bioreactor is configured to operate (a) in the operating parameter regime of the hydrogen-rich product stream described herein or (b) in the operating parameter regime of the high-reducing product to the hydrogen product stream described herein.

[0183] In various embodiments, the oxygen produced by the electrolyzer is used in an integrated process such as a gasification process. For example, a system having a Fischer-Tropsch reactor can use an electrolyzer configured to receive a carbon dioxide input from the partial oxidation of hydrocarbons that consume oxygen or from the gasification of biomass. The oxygen may come from the anode side of the carbon dioxide electrolyzer and / or from an air separation unit.

[0184] FIG. 5 shows a system 501 that includes a carbon dioxide electrolyzer 503 configured to produce syngas. System 501 also includes an aerobic fermentation reactor 505 configured to produce syngas. System 501 is also configured to deliver oxygen produced at the anode of electrolyzer 503 to reactor 505. This oxygen can replace a portion of the oxygen normally provided from alternative sources such as air separation, thus reducing the energy duty and energy scale of the air separation unit.

[0185] System 501 is configured to deliver syngas produced by electrolyzer 503, fermentation reactor 505, and potentially other sources to a downstream gasification unit or partial oxidation unit 507. System 501 may be configured to deliver some waste carbon dioxide from unit 507 to the cathode input stream of electrolyzer 503. Naphtha and Fuel

[0186] As indicated, the Fischer-Tropsch reaction can be characterized by the following general formula. (2n + 1)H2 + nCO → CnH2n + nH2O

[0187] The following discussion focuses on the Fischer-Tropsch reaction, but one of ordinary skill in the art will understand that related reactions of a particular class may be used to produce liquid hydrocarbons and mixtures thereof (often commonly referred to as naphtha) from an input stream containing hydrogen and carbon monoxide. The class of reactions produces various compositions of liquid hydrocarbon mixtures depending on the composition of the input stream and the reaction conditions. Although the term Fischer-Tropsch is used herein, it should be understood to encompass any of the classes of reactions that produce naphtha from a mixture containing carbon monoxide and hydrogen. Generally, such reactions are exothermic.

[0188] In various embodiments, the input stream to the Fischer-Tropsch reactor is a CO:H2 molar ratio of about 1:2. To use CO2 as a starting point for producing a CO / H2 mixture (or other Fischer-Tropsch input), some conventional non-electrolytic processes require two steps. For example, a conventional process uses a first process to produce CO2 + H2 (step 1), and then uses the reverse water gas shift (RWSG) reaction (step 2) to react CO2 + H2 to produce CO and water, resulting in a gas having a ratio close to the required 2:1 CO:H2. Thus, in a conventional process, the Fischer-Tropsch reaction can be used to produce liquid hydrocarbons only after obtaining CO and hydrogen in the correct ratio. Water gas shift (WSG) reaction and reverse water gas shift reaction catalysts can produce metal dust that is harmful to downstream processes. Further, the water gas shift reaction requires a supply of carbon monoxide and / or hydrogen.

[0189] It should be noted that conventional syngas processes may be used to directly produce a CO + H2 mixture (rather than using a WSG reaction and / or RWSG reaction or a carbon dioxide electrolyzer that may emphasize the production of CO). However, syngas production often uses coal.

[0190] A Fischer-Tropsch system that uses a carbon dioxide electrolyzer as a carbon monoxide source has various advantages over the WSG or syngas routes. For example, unlike the RWSG reaction, the carbon dioxide electrolyzer does not produce metal dust. Furthermore, compared to the RWGS reaction, the carbon dioxide electrolyzer achieves a higher conversion of CO2 to CO.

[0191] However, the carbon dioxide electrolyzer may not produce a gas having a molar ratio of CO:H2 of approximately 1:2 required for Fischer-Tropsch feed. In some cases, the carbon dioxide electrolyzer produces a CO-rich stream. Thus, in some embodiments, a Fischer-Tropsch system, or any other system that requires a mixture of carbon monoxide and hydrogen, can use a water electrolyzer or any other hydrogen source that functions in combination with, optionally, a carbon dioxide electrolyzer. The water electrolyzer is configured to produce gaseous hydrogen to supplement the CO-rich output of the carbon dioxide electrolyzer. In some embodiments, a relatively hydrogen-rich syngas can be produced as part of the co-electrolysis of carbon dioxide and water. To achieve an approximate 1:2 CO:H2 feed concentration for the F-T reaction, the system can include sensors configured to determine the concentrations of CO and H2 passing through a gas separation unit from the CO2 electrolyzer. The detected information can be used as feedback to adjust the operating conditions of the water electrolyzer to deliver a hydrogen stream having the amount of H2 necessary to bring the overall stream to an approximately 1:2 CO:H2 concentration.

[0192] Alternatively, a single CO2 electrolyzer can be used to produce a suitable feed mixture of Fischer-Tropsch CO and H2. This can be achieved by operating the electrolyzer in such a way as to bias the output towards hydrogen production and / or by treating the output of the electrolyzer to adjust its composition and then delivering it to a Fischer-Tropsch reactor. In certain embodiments, the carbon dioxide electrolyzer includes a membrane electrode assembly (MEA) that allows a relatively high proportion of H+ to reach the cathode. One way to promote a relatively high H+ flux at the cathode is to use a bipolar MEA with a relatively thin cathode buffer layer and / or to use a cathode and cathode buffer layer with a polymer having a relatively high H+ transport rate. In another approach, the carbon dioxide electrolyzer is configured or operated to deplete carbon dioxide. In certain embodiments, the electrolyzer is operated at a relatively high current density and tends to have a high ratio of hydrogen to carbon monoxide. In some implementations, the electrolyzer uses both a relatively high current density to the electrolyzer and a relatively low carbon dioxide supply. Operating at a relatively high current density has the advantage of being produced using a relatively inexpensive electrolyzer due to the cost of the equipment.

[0193] The output of the CO2 electrolyzer contains product CO, by-product H2, unreacted CO2, and water vapor. The system may be configured to remove water vapor and separate unreacted carbon dioxide. A gas separation unit can be used to separate CO2 from CO and H2 and / or to concentrate CO and H2. The system can include a recirculation loop for recirculating water to the water inlet of the CO2 or water electrolyzer. The separated unreacted CO2 is then compressed and returned to the inlet of the CO2 electrolyzer. Examples of gas separation units are shown in FIGS. 19, 20, 23A-23D, and the associated description.

[0194] The F-T reactor can operate at pressures above about 300 psi and temperatures of about 150 - 300 °C. If the outputs of the carbon dioxide electrolyzer and any water electrolyzer are not at the required pressure, the system can use a compressor to raise the feed gas pressure before entering the F-T reactor. Inside the F-T reactor, the CO-H2 mixture is converted into raw F-T liquids and waxes. The system can include a separator following the F-T reactor to separate water, high melting point F-T liquids, medium melting point F-T liquids, and a tail gas, a mixture of volatile hydrocarbons, CO2, CO, and H2. The F-T liquids can be further upgraded via hydrocracking. Distillation and separation of different fractions of the F-T liquids can result in jet fuel, diesel fuel, and gasoline. The water from the F-T reactor can be filtered to remove impurities and supplied to the water input of the CO2 and / or any water electrolyzer.

[0195] The F-T system can be designed such that the tail gas and / or volatile hydrocarbons (e.g., including methane) are recycled back to the CO2 electrolyzer. The system may be configured to separate the tail gas into CO2, which can be compressed and supplied directly to the electrolyzer inlet and the volatile hydrocarbons and unreacted CO and H2. The system may be designed or configured such that these products are supplied to a combustion reactor to produce heat, energy, and CO2. The CO2 is then supplied to the CO2 electrolyzer inlet. The O2 from the electrolyzer may be used as an oxygen source for combustion, resulting in a pure CO2 output stream. The combustion reactor may be operated in a "rich burn" mode using an excess of fuel relative to oxygen to minimize the oxygen concentration in the outlet stream. The water from the combustion reaction may be separated from the gas output and supplied to the water input of the CO2 electrolyzer or the water electrolyzer.

[0196] Since the Fischer-Tropsch reaction is exothermic, it generates heat that can be used for other purposes within the system. Examples of such other uses include separation (e.g., distillation of light hydrocarbons) and reactions. In conventional systems, such reactions are endothermic reactions for the production of synthesis gas, such as reforming of fossil fuels, gasification of biomass, or production from carbon dioxide and hydrogen by the reverse water-gas shift. Thus, in conventional processes, all or a substantial portion of the excess heat from the Fischer-Tropsch reaction is typically directed towards synthesis gas production. However, in this case where synthesis gas is produced at low temperatures (e.g., less than about 100 °C) by processes such as carbon dioxide electrolysis optionally together with low-temperature water electrolysis, there is more excess heat from the Fischer-Tropsch reaction available for other processes such as carbon dioxide capture, thereby reducing the overall external heat requirement of the system and improving the carbon and energy efficiency of carbon dioxide for the fuel synthesis route.

[0197] In some embodiments, the tail gas is fed to a reformer where methane or other gaseous hydrocarbons react with water to produce a mixture of hydrogen and carbon monoxide, which is a form of synthesis gas. This can increase the yield of carbon from carbon dioxide in the liquid hydrocarbon product. Depending on the composition of the tail gas, the ratio of hydrogen to carbon monoxide can vary. In some embodiments, some amount of carbon dioxide and / or oxygen is present in the reformer. Often, the reforming reaction is endothermic. In some embodiments, the heat to drive the endothermic reaction is provided at least in part from the excess heat generated during the Fischer-Tropsch reaction. In some cases, some heat may also be provided by combustion or direct electrical heat. In the case of heat from combustion, oxygen may optionally be fed to the furnace (from the electrolyzer) to improve efficiency, and carbon dioxide emissions can be captured and fed to the electrolyzer.

[0198] Figure 6A shows a system 601 configured to produce liquid hydrocarbons that are carbon dioxide feedstocks such as those in which the primary or dedicated source of carbon contains carbon dioxide and / or carbon monoxide. The system includes two primary reactors, namely an electrolytic carbon dioxide reduction cell or electrolyzer 611 and a Fischer-Tropsch reactor 621.

[0199] The electrolyzer 611 is connected to a power source and has one or more inlets for receiving reactants such as carbon dioxide and water. The electrolyzer 611 has one or more anodic outlets for removing oxygen and, optionally, trace impurities, and one or more cathodic outlets for removing reduction products containing at least carbon monoxide. Other compounds exiting the cathodic side may include hydrogen, water, and carbon dioxide.

[0200] The cathodic outlet is connected to a purification unit such as a carbon monoxide purification unit 612 designed to separate or purify carbon monoxide from other components. In the illustrated embodiment, the purification unit 612 has one outlet for supplying carbon monoxide and another outlet for supplying carbon dioxide, hydrogen, and, optionally, some carbon monoxide. In certain embodiments, the carbon monoxide purification unit 612 may be an adsorbent-based unit as shown in FIGS. 19, 20, and / or FIGS. 23A-23D and the associated description.

[0201] In system 601, carbon dioxide is recycled back from the outlet of the CO purification unit 612 to the inlet stream on the cathodic side of the electrolyzer 611, optionally with some hydrogen and carbon monoxide.

[0202] The Fischer-Tropsch reactor 621 is configured to receive carbon monoxide and hydrogen in a pressurized feed stream and in a specified composition. In system 601, compressor 624 compresses the carbon monoxide from electrolyzer 611, along with hydrogen, to an appropriate pressure for the Fischer-Tropsch reaction. The Fischer-Tropsch reaction can occur at a temperature of about 150 - 300 °C and a pressure of about 1 to several tens of atmospheres. Since the reaction is exothermic, little or no heat is provided to reactor 621.

[0203] As described, the input to the Fischer-Tropsch reactor can have a CO:H ratio of about n:(2n + 1), where n is the length of the carbon atoms of the desired alkane product of the reaction. Thus, in various embodiments, the molar ratio of hydrogen to carbon monoxide provided to reactor 621 is about (2n + 1) to n. To provide the desired inlet composition ratio of hydrogen to carbon monoxide for the Fischer-Tropsch reaction, hydrogen source 614 may be coupled to the outlet of CO purification unit 612 or the inlet of compressor 624. Alternatively, or in addition, electrolyzer 611 may be designed or operated such that the ratio of hydrogen to carbon dioxide is relatively high. The reactor design and operating conditions for achieving this ratio are described elsewhere in this specification. In some cases, a gas with a relatively high ratio of hydrogen to carbon monoxide is produced from a reforming reaction such as a reaction using FT tail gas as an input.

[0204] As shown, system 601 is configured to provide the output of Fischer-Tropsch reactor 621 to a separator 623 configured to separate MFTL and HFTL Fischer-Tropsch liquids from water and tail gas. As shown, Fischer-Tropsch water may be recycled back to the input of CO purification unit 612 and / or the input of electrolyzer 611.

[0205] System 601 comprises a main recirculation loop having a separation unit 631, a combustion chamber 632, and a water / gas separator 633. The separation unit 631 receives tail gas from the separator 623 and is configured to remove carbon dioxide from the volatile hydrocarbons. System 601 is configured to recycle carbon dioxide from the unit 631 to the carbon dioxide feed stream to the electrolyzer 611.

[0206] System 601 is configured to transport volatile hydrocarbons from the separation unit 631 to the combustion unit 632, and the combustion unit is configured to combust the hydrocarbons using an oxygen source from the electrolyzer 611. System 601 is configured to transport the combustion products from the combustion unit 632 to the gas / water separator unit 633, and the gas / water separator unit is configured to separate the combustion products of carbon dioxide and water. System 601 is configured to transport water to the anode inlet of the electrolyzer 611 and carbon dioxide to the cathode inlet of the electrolyzer 611.

[0207] In certain embodiments, a carbon dioxide electrolyzer disposed upstream from the Fischer-Tropsch reactor is configured to operate in (a) the operating parameter regime of the hydrogen-rich product stream described herein, and / or (b) the operating parameter regime of the high reduction product to CO2 ratio described herein.

[0208] In certain embodiments, System 601 includes one or more carbon dioxide capture units that include an adsorbent for capturing carbon dioxide during a first stage and releasing carbon dioxide during a second stage. The separation unit 631 and / or the gas / water separator unit 633 may include or be configured to cooperate with such carbon dioxide capture units. By way of example, several operating principles are provided in the description of the direct air capture unit described herein. In some embodiments, the Fischer-Tropsch system is configured to provide waste heat generated from the exothermic Fischer-Tropsch reaction to the carbon dioxide capture unit.

[0209] Figure 6B presents an exemplary system 634 for generating a liquid hydrocarbon mixture from a carbon dioxide input stream 635 by using (a) a carbon dioxide electrolyzer 636 for generating carbon monoxide and hydrogen 637 and (b) a Fischer-Tropsch reactor 638 configured to receive the carbon monoxide and hydrogen and generate liquid hydrocarbons. The carbon monoxide and hydrogen are at least partially generated by the electrolyzer 636 and are pretreated in a syngas processing element 640 that can purify the syngas before delivery or otherwise reform it (e.g., removal of unreacted CO2 from the electrolyzer, as well as compression and / or heating or cooling of the syngas stream) before entering the Fischer-Tropsch reactor. The system 634 is further configured to provide the processed gas from the element 640 to the Fischer-Tropsch reactor 638, which can generate mixed light hydrocarbons and other components 642. The system enables a product separation subsystem 643 to be available, which can include a mechanism for separating tail gas 641 from one or more liquid hydrocarbon streams 644. In the illustrated embodiment, the system 634 includes a reformer 645 configured to receive the tail gas 641. The tail gas includes methane that can react with water (optionally also included in the tail gas 641) by a methane reforming reaction to produce a hydrogen-rich mixture 647 of carbon monoxide and hydrogen. The system 634 is also configured to deliver the mixture 647 to the syngas processing element 640, which prepares the gas for introduction into the Fischer-Tropsch reactor 638. The methane reforming reaction is endothermic. In some embodiments, excess heat from the reaction in the Fischer-Tropsch reactor 638 is provided to the reformer 645.

[0210] In the embodiment shown in FIG. 6B, the system 634 is optionally configured to provide oxygen 649 from the electrolyzer 636 to a furnace 651, which is configured to burn fuel and generate additional heat for use in the system 634 or elsewhere. Direct air capture of CO2

[0211] In certain embodiments, the electrolytic carbon dioxide reduction system uses carbon dioxide received directly from the air. A system for such embodiments includes a direct air CO2 capture subsystem and a carbon dioxide reduction electrolyzer subsystem. The system is configured such that CO2 from the capture subsystem supplies CO2 directly or indirectly to the cathode side of the electrolyzer subsystem.

[0212] Since air is often the only significant feedstock, an air-captured CO2 electrolysis system may be located anywhere there is space for the system components. In some arrangements, the system occupies relatively unpopulated areas. In some arrangements, the system occupies residential areas. In some embodiments, the system is located at least in part on a vehicle or a ship. For example, the air capture unit may be provided on a vehicle or a ship, and the carbon dioxide electrolyzer may be provided at a port or an offshore platform. In some cases, the location is ready for an energy supply, such as a location rich in sunlight and / or wind. In some cases, the location is a desert. In some embodiments, the system is located in an extraterrestrial environment having a CO2-containing atmosphere. In some embodiments, the system is located on a large ship such as a cargo ship or a military ship such as an aircraft carrier. In some embodiments, the energy source is provided by a solar power plant or a wind farm associated with an offshore platform or a port, and the carbon dioxide capture unit is provided on a ship or another vessel. The carbon dioxide electrolyzer may be provided on an offshore platform or a port.

[0213] The system may be designed such that air or other gas is provided to the CO2 capture subsystem under specified conditions. In certain embodiments, a fan, a vacuum pump, or simply wind is used to deliver air to the CO2 capture subsystem.

[0214] In certain embodiments, the CO2 capture subsystem includes two stages, namely a first stage (stage 1) in which an adsorbent that removes CO2 from air is brought into contact with air, and a second stage (stage 2) in which heat, electricity, pressure, and / or humidity are applied to the adsorbent to release CO2 and / or water.

[0215] In some implementations, the CO2 capture subsystem uses a solid or liquid absorbent or adsorbent to capture CO2 in stage 1. In various implementations, stage 1 is carried out at or near ambient conditions. In stage 2, a temperature, electrical, pressure, and / or moisture swing is applied, and the absorbed or adsorbed CO2, and optionally water, are released.

[0216] In certain embodiments, the absorbent is heated to release CO2. As an example, the adsorbent is heated from an ambient temperature (e.g., about 20 - 40 °C) to at least about 75 °C to release CO2 and optionally water. In some cases, the temperature variation is from ambient to about 50 - 1000 °C, or from ambient to about 75 - 200 °C, or from ambient to about 600 - 1000 °C. As an example, the adsorbent is heated for a time sufficient to remove the desired proportion of CO2, and optionally water. The duration is a function of the amount of adsorbent being processed, the proportion of CO2 and / or water being removed, and the heat transfer to the adsorbent.

[0217] In certain embodiments, the absorbent is exposed to moisture to release CO2. As an example, the adsorbent is first exposed to dry air (e.g., air having less than about 50 mol% water, or less than about 30 mol% water, or less than about 5 mol% water), and subsequently exposed to moist steam (e.g., air having at least about 75 mol% water, or at least about 90 mol% water, or about 100 mol% water).

[0218] In some embodiments, the CO2 capture unit uses an electro-swing mechanism to capture and then release CO2. In certain cases, the electro-swing carbon dioxide unit includes a Faradaic adsorption system that includes an electrochemical cell that utilizes the reductive addition of CO2 to a redox species such as a quinone (e.g., 2,6-di-tert-butyl-1,4-benzoquinone), 4,4'-bipyridine, or a thiolate for carbon dioxide capture. These redox agents may be provided in an organic electrolyte. In some cases, the electro-swing adsorption system provides a carbon dioxide capture material on a solid support such as a carbon nanotube support and / or a zeolite support. In some cases, the electro-swing CO2 capture unit releases CO2 by supplying heat (e.g., by Joule heating) to an absorbent and / or an electrode that holds the captured CO2.

[0219] Depending on the configuration of the CO2 capture subsystem and its operating conditions, CO2 can be produced from air at a high concentration, e.g., about 90 mol% or more. In some cases, the CO2 capture subsystem is configured to produce CO2 at a relatively low concentration that is sufficient for the CO2 reduction electrolyzer to operate.

[0220] As an example, CO2 capture adsorbents and related subsystem components are available from Climeworks AG, Zurich, Switzerland; Global Thermostat, New York, NY; Carbon Engineering Ltd., Squamish, B.C., Canada; and Silicon Kingdom Holdings, Dublin, Ireland.

[0221] As indicated, the captured and subsequently released CO2 is a feedstock that is delivered directly or indirectly to the cathode side of the CO2 reduction electrolyzer. In certain embodiments, water captured from air is also used as a feedstock for the CO2 electrolyzer.

[0222] In certain embodiments, the air capture CO2 electrolysis system is configured to operate to deliver CO2 from the direct air capture subsystem in a substantially pure stream of, for example, about 99 mole % or more CO2. In certain embodiments, the system is configured to operate using a low concentration of CO2 to the electrolyzer, such as about 98 mole % or more CO2, or about 90 mole % or more CO2, or about 50 mole % or more CO2. In some cases, very low CO2 concentrations are used as feedstock. Such concentrations are still substantially higher than the atmospheric concentration of carbon dioxide, which is about 0.035 mole %. In certain embodiments, the system is configured to operate using a CO2 concentration of about 5 - 15 mole % that is mixed with another gas such as air or nitrogen.

[0223] Depending on the type of adsorbent used in the process, water may also be captured with the CO2 and released with the CO2. In certain embodiments, the output of the CO2 capture subsystem is humidified CO2 having a water concentration of about 0 to 20 mole % water.

[0224] In certain embodiments, the output of the CO2 capture subsystem contains only CO2 and other components in air such as nitrogen, oxygen, water, argon, or any combination thereof. In all cases, the CO2 is present at a higher concentration than in air. In certain embodiments, the output of the CO2 capture subsystem is sulfur - free.

[0225] The direct air capture unit and the CO2 electrolyzer can be integrated in several ways depending on the type of air capture technology. The heat and mass transfer components may be incorporated throughout the air capture CO2 electrolysis system.

[0226] For example, in some designs, the CO2 reduction electrolyzer is configured to receive CO2 directly from the air capture subsystem and provide heat and / or humidity directly to the air capture subsystem. The heat provided can release the CO2 captured during stage 2 of the air capture subsystem that uses a temperature swing desorption mechanism. The humidified electrolyzer product gas can be used to release the CO2 captured during stage 2 of the air capture subsystem that uses a moisture swing desorption mechanism.

[0227] In certain embodiments, the CO2 electrolyzer is designed or configured to receive diluted CO2 (e.g., CO2 at about 50 mol% or less) as an input.

[0228] The direct air capture unit can be designed using a plurality of adsorbent vessels. To receive a continuous flow of CO2 (and optionally water) from the air capture subsystem, at least two different vessels operate such that they are in different stages of sorption / desorption during the operation of the entire air capture CO2 electrolysis system. For example, while one adsorbent vessel is taking in air to capture CO2, another adsorbent vessel may be heated to release CO2, and as each vessel continues its sorption / desorption cycle, the sorption vessel that was taking in CO2 releases CO2 and vice versa. By adding more vessels at different points in the cycle, a continuous input stream can be delivered to the CO2 electrolyzer and a continuous flow of air containing CO2 as well as moisture and / or heat and / or vacuum can be received.

[0229] The direct air capture unit can be sized to deliver a desired amount of CO2 flow for a downstream process such as a CO2 electrolyzer. This can include using vessels containing multiple adsorbents. For example, the direct air capture subsystem may be configured to deliver 750 slpm of CO2. Such a subsystem can operate at 300 mA / cm 2 and 3 V / cell with 1000 cm 2Assuming that when combined with a 200-cell electrochemical stack consisting of a membrane-electrode assembly, 90% of the CO₂ is provided for the CO current efficiency of the process, 378 slpm of CO and 42 slpm of hydrogen can be produced. Unreacted CO₂ at the outlet of the electrolyzer may be recycled to the inlet to increase carbon efficiency. When operated continuously, a unit combining air capture and an electrolyzer can produce approximately 675 kg / day of CO. Generally, in some designs, an air capture CO₂ electrolyzer system is configured to output at least about 100 kg / day of CO and / or other CO₂ reduction products, and in some designs, an air capture CO₂ electrolyzer system is configured to output at least about 500 kg / day of CO and / or other CO₂ reduction products.

[0230] In certain embodiments, a system that uses a carbon dioxide electrolyzer and optionally direct air capture of a carbon dioxide unit also includes a module configured to capture water from air or the atmosphere. In some embodiments, the module configured to capture water from air utilizes solar energy from a photovoltaic and / or a thermal solar with a hygroscopic material. In certain embodiments, the module configured to capture water is an ambient dehumidifier such as a hydro panel (e.g., available from 0mAss Water, Inc. of Scottsdale, Arizona).

[0231] FIG. 7A shows an air capture CO₂ electrolyzer system 701 comprising a direct air CO₂ capture subsystem 703 and a CO₂ reduction electrolyzer subsystem 705. As shown, the direct air CO₂ capture subsystem 703 is configured to receive air containing CO₂, optionally with humidity, for example under atmospheric conditions (about 0.035 mol% CO₂), during the sorption stage 1, and release air with most of the CO₂ removed and optionally much of the humidity removed.

[0232] The direct air CO2 capture subsystem 703 is configured to release CO2 and optionally water during stage 2. At least CO2, and optionally water, are provided as an input to the CO2 electrolyzer 705. The CO2 released from the direct air capture subsystem 703 during stage 2 is provided to the cathode side of the electrolyzer 705. As shown, an optional CO2 purification unit 707 is inserted between the direct air CO2 capture subsystem 703 and the electrolyzer 705. The water optionally provided by the direct air CO2 capture subsystem 703 may be directed to the cathode side (as humidity in the CO2 feedstock) or the anode side (as a reactant) of the electrolyzer 705.

[0233] In the illustrated embodiment, the electrolyzer 705 is configured to receive electricity (to facilitate the CO2 reduction reaction and the anode oxidation reaction). The electrolyzer 705 is also configured to provide excess heat from the electrolysis reaction to facilitate the air CO2 capture subsystem 703 and drive stage 2 (CO2 release from the adsorbent).

[0234] The CO2 electrolyzer 705 is configured to output oxygen (anode reaction product if water is a reactant) and one or more CO2 reduction products that may include CO and / or other carbonaceous products as described elsewhere herein. The product stream of the CO2 electrolyzer 705 may include hydrogen, CO2, and / or water. As shown, the system 701 is configured to provide the electrolyzer output to a separation unit 709 configured to separate CO and / or other carbonaceous electrolysis products from hydrogen, CO2, water, and / or other components. In the illustrated embodiment, the system 701 is configured to deliver humidified CO2 from the separation unit 709 and direct it to the direct air CO2 capture subsystem 703.

[0235] In certain embodiments, when fuel is burned, carbon dioxide is captured within a ship or vehicle. The fuel can be used, for example, in an internal combustion engine to propel a ship or vehicle. The fuel may also be used for other purposes such as heating, power generation, etc. The captured carbon dioxide is provided to the cathode of a carbon dioxide electrolyzer and produces reduction products that are directly used, stored, or converted to different products (e.g., chemicals, polymers, or fuels) by downstream processing. In embodiments where the electrolysis device and related downstream components are configured to produce fuel, the resulting fuel can be used in the original ship or vehicle, or in one or more other ships or vehicles. In some implementations, both the carbon dioxide capture subsystem and the electrolyzer, as well as any downstream processing subsystems, are installed on a vehicle or ship. In some implementations, only the carbon dioxide capture subsystem is provided on a vehicle or ship. In some cases, the captured carbon dioxide is temporarily stored on the vehicle or ship. For example, carbon dioxide can be stored in one or more tanks, pressure vessels, tanker ships, etc. In other cases, the captured carbon dioxide is stored away from the vehicle or ship, such as in an underground tank, tanker ship, offshore platform, etc. In some cases, the stored carbon dioxide is unloaded from the ship or vehicle and provided to the carbon dioxide electrolyzer there. Examples of locations where the stored carbon dioxide may be unloaded and / or where the electrolyzer is located include chemical plants, ports, and offshore platforms that are located close to green energy sources such as wind energy or solar energy. Examples of ships and vehicles include ships, trucks, buses, passenger cars, aircraft, and other ships.

[0236] Figure 7B is a diagram showing an example in which carbon dioxide is captured from fuel combustion products mounted on a ship or vehicle. As shown, ship 720 produces carbon dioxide 721, which is stored in storage medium 722 before being supplied to carbon dioxide electrolyzer 723. The electrolyzer and related downstream chemical processing equipment can produce chemical product 725 and / or fuel 726. In the case of fuel, the fuel may be utilized by the vehicle or ship.

[0237] In certain embodiments, depending on the needs of the system, the carbon dioxide electrolyzer disposed downstream of the direct air CO2 capture subsystem operates in (a) the operating parameter regime of high reduction product to CO2 ratio described herein, (b) the operating parameter regime of hydrogen-rich product stream described herein, or (c) the operating parameter regime of high reduction product to hydrogen product stream described herein. Polycarbonate

[0238] Certain aspects of the present disclosure relate to a polycarbonate production system that includes (a) one or more carbon dioxide electrolyzers configured to produce one or more carbon-containing products and (b) one or more polycarbonate synthesis reactors configured to produce a polycarbonate polymer from a carbon-containing compound extracted directly or indirectly from the products of the one or more electrolyzers.

[0239] In certain embodiments, at least one electrolyzer in the polycarbonate production system includes a membrane electrode assembly (MEA) optionally including a polymer electrolyte membrane (PEM), such as a cation exchange polymer membrane. Unless otherwise specified or apparent from the context of this specification, a carbon dioxide reduction electrolyzer, including a carbon dioxide electrolyzer, encompasses an MEA-based electrolyzer, and specific embodiments thereof are described elsewhere in this specification.

[0240] In certain embodiments, the polycarbonate production system includes a carbon dioxide reduction electrolyzer configured to produce carbon monoxide and one or more other subsystems configured to convert the carbon monoxide to one or more intermediates for subsequent reaction to produce a polycarbonate polymer.

[0241] In various embodiments, the direct output of the carbon dioxide electrolyzer is converted to one or more intermediate compounds such as phenol, ketones, and / or organic carbonates, which react to produce polycarbonates. The conversion to such intermediate compounds can occur by any of a variety of processes. Examples include the Fischer-Tropsch reaction, gas fermentation reaction, and decomposition reaction.

[0242] In an embodiment of the gas fermentation subsystem, a carbon dioxide electrolyzer is used to produce carbon monoxide, and optionally hydrogen, which is then used in a downstream gas fermentation process to produce one or more intermediate compounds for producing a polycarbonate polymer. Examples of these intermediate compounds include, for example, ketones (e.g., acetone), light hydrocarbons, phenols, and the like.

[0243] In a Fischer-Tropsch subsystem that may correspond to other embodiments described herein, carbon monoxide and hydrogen from the carbon dioxide electrolyzer are reacted to form naphtha or other light hydrocarbon products.

[0244] In certain embodiments, one or more subsystems for producing intermediates are configured to produce or produce a diol compound such as bisphenol A. In another embodiment, other diol polycarbonate intermediates may be produced, which include bisphenols other than bisphenol A. In certain embodiments, the subsystem includes a reactor for producing phosgene from carbon monoxide and chlorine. Carbon monoxide may be produced from a carbon dioxide reduction electrolyzer.

[0245] In certain embodiments, the carbonate synthesis system uses at least two separate electrolysis modules that can share some common electrical infrastructure, such as a common electrical bus. In certain embodiments, the two separate electrolysis modules are two separate electrolytic carbon dioxide reduction cells. As an example, the first carbon dioxide reduction cell is a carbon dioxide electrolysis cell designed, configured, or operated to produce carbon monoxide and optionally hydrogen gas, and the second carbon dioxide reduction cell is a carbon dioxide electrolysis cell designed, configured, or operated to produce at least one product compound having at least two carbon atoms, such as a ketone such as ethylene or acetone. In some implementations, the first electrolysis cell comprises a cathode having a noble metal catalyst such as gold, and the second electrolysis cell comprises a cathode having a transition metal catalyst such as copper. As disclosed elsewhere herein, MEA-based carbon dioxide electrolysis cells can have various designs or configurations that enable the production of distinct products (e.g., CO vs. C2 compounds).

[0246] In certain embodiments, the first electrolysis module is a carbon dioxide reduction module, and the second electrolysis module is a chlorine generation module such as a chlor-alkali cell. These two modules can share a common electrical infrastructure. In some embodiments, the chlorine generation module is a conventional chlor-alkali module configured to receive a chloride salt and water as inputs and generate chlorine gas and hydrogen gas as outputs. In some embodiments, the chlorine generation module includes an oxygen reduction chlor-alkali cell configured to receive a chloride salt in an electrolyte, receive oxygen gas at a cathode (oxygen depolarized cathode), generate chlorine gas at an anode, and generate water at the cathode. Such oxygen reduction chlor-alkali cells operate more efficiently than conventional chlor-alkali cells and consume less electrical energy. However, they require an oxygen source. In some implementations, the polycarbonate production system is configured such that the oxygen gas generated as the anode of the carbon dioxide reduction cell is provided to the cathode of the chlorine generation cell configured to reduce oxygen. In some implementations, the polycarbonate production system is configured to provide the water generated by the oxygen reduction chlor-alkali cell to the carbon dioxide removal cell either as anode water or for humidification of the carbon dioxide feed stream.

[0247] In a particular embodiment shown in FIG. 8A, system 801 includes an electrolysis cell such as a chlor-alkali cell for generating chlorine. In some implementations, system 801 is configured to supply oxygen from electrolyzer 803 to the cathode of a chlor-alkali cell including an oxygen depolarized cathode. In some implementations, system 801 is configured to supply oxygen from a water electrolyzer to the oxygen depolarized cathode of a chlor-alkali cell. System 801 may be configured to provide hydrogen from a water electrolyzer to a Fischer-Tropsch reactor, a gas fermentation reactor, or other reactors used in the production of polyol precursors. Instead of oxygen from other sources such as air separation, oxygen from electrolyzer 803 or other electrolyzers may be used.

[0248] In certain embodiments, the alkali by-products of the chlor-alkali cell (e.g., sodium hydroxide) are provided as feedstock to complementary chemical production systems such as, by way of example, a formate production system that uses an oxidation carbon electrolyzer.

[0249] In certain embodiments, the polycarbonate uses a bisphenol A linkage in the polymer backbone, and indeed, examples of most of the systems and methods presented herein describe bisphenol A as a polycarbonate precursor, along with phosgene. However, in some applications, other diols are used in place of bisphenol A. Examples include other linear and cyclic unsaturated diols, as well as diphenols and other bisphenols. It should be understood that when referring to bisphenol A in the examples described herein, it is intended that other bisphenols may be appropriately used for the desired polycarbonate end product. Also, appropriate changes to the systems and methods may be employed to replace the phenol production module with a module configured to produce phenol derivatives or analogs and / or to replace the acetone production module with a module configured to produce other ketones.

[0250] The polycarbonate synthesis reaction may include the treatment of bisphenol A with sodium hydroxide that deprotonates the hydroxyl groups of bisphenol A. (HOC6H4)2CMe2 + 2NaOH → Na2(OC6H4)2CMe2 + 2H2O

[0251] The diphenoxide (Na2(OC6H4)2CMe2) reacts with phosgene to give a chloroformate, which is then attacked by another phenoxide. The net reaction from the diphenoxide is Na2(OC6H4)2CMe2 + COCl2 → 1 / n[OC(OC6H4)2CMe2] n + 2NaCl.

[0252] In various embodiments, at least one carbon dioxide reduction electrolyzer used in a polycarbonate production system is designed or configured to operate to produce a significant proportion of hydrogen in addition to carbon monoxide. In certain embodiments, a carbon dioxide electrolyzer disposed upstream of a diol production reactor is configured to operate in (a) the operating parameter regime of the hydrogen-rich product stream described herein, and / or (b) the operating parameter regime of the high reduction product to CO2 ratio described herein.

[0253] Electrolyzers for producing carbon monoxide often use a cathode catalyst containing a noble metal such as gold. Such catalysts are more favorable for the production of carbon monoxide than for the production of hydrogen-containing compounds such as methane, ethylene, and formic acid. An electrolyzer configured to provide a hydrogen-rich product can use a design that (a) depletes the cathode of the carbon dioxide reactant, and / or (b) allows for a relatively high flux of hydrogen ions to be transported from the anode where they are generated to the cathode. An electrolyzer that can operate with a relatively low input of carbon dioxide has a flow field or gas diffusion component that limits the access of carbon dioxide to the active sites of the electrolyzer cathode. An electrolyzer that can operate with a relatively high hydrogen ion flux to the cathode can have a membrane electrode assembly (MEA) with a cation-conductive polymer and / or a mixed ion-conductive polymer at the cathode, and / or can have a cathode buffer layer if one is used. If the MEA includes an anion-conductive cathode buffer layer, the layer is designed to have a relatively poor hydrogen ion transport rate and / or a relatively high hydrogen ion transport rate.

[0254] Figure 8A shows an overall view of a system for producing a polycarbonate polymer using a carbon dioxide reduction electrolytic cell. As shown, the polycarbonate production system 801 includes a carbon dioxide reduction electrolytic cell 803 configured to promote an anodic reaction and a cathodic reaction that receive carbon dioxide and water as reactants and electricity and produce oxygen and one or more carbon dioxide reduction products. In the illustrated embodiment, the carbon dioxide reduction reactor 803 is configured to produce at least carbon monoxide as a reduction product. The system 801 is specifically configured to deliver carbon monoxide from the electrolytic cell 803 to the phosgene production reactor 805. The reactor 805 further includes an input for receiving chlorine gas. Chlorine gas and carbon monoxide react within the phosgene reactor 805 to produce phosgene as an output. The system 801 is further configured to deliver phosgene from the phosgene reactor 805 to the polycarbonate synthesis reactor 807.

[0255] In certain embodiments, the carbon dioxide electrolytic cell disposed upstream from the phosgene production reactor is configured to operate in (a) the operating parameter regime of the high reduction product to hydrogen production stream described herein, and / or (b) the operating parameter regime of the high reduction product to CO2 ratio described herein.

[0256] As shown, polycarbonate synthesis reactor 807 also has an input for receiving a diol input material. The input of the diol can be generated by various methods including via a reactor not shown in this figure. Alternatively, as depicted herein, system 801 includes a diol synthesis reactor or subsystem 809 configured to receive a carbon dioxide reduction product from electrolyzer 803. In various embodiments, these electrolyzer reaction products include carbon monoxide and hydrogen. In some cases, these electrolyzer products include C2 or higher products such as acetone or formaldehyde. In certain embodiments, reactor or subsystem 809 is configured to receive input from sources other than electrolyzer 803 and react. These other inputs may include, for example, phenolic compounds such as bisphenol. System 801 is also configured to transport the diol produced by diol synthesis reactor or subsystem 809 to polycarbonate synthesis reactor 807. Within polycarbonate synthesis reactor 807, the diol reacts with phosgene to produce a polycarbonate polymer. In the illustrated embodiment, the polycarbonate final product is available from polycarbonate synthesis reactor 807 via an outlet.

[0257] It should be understood that a polycarbonate production system such as that shown in FIG. 8A may include additional types or alternative types of modules not shown in the figure. These include, for example, one or more purification units such as a carbon monoxide purification module, a heater, a compressor, a condenser, and other chemical reactors. Examples of gas purification units for use in the system of 8A or any other polycarbonate production system described herein are presented in FIGS. 19 and 20 and the associated description.

[0258] The types of reactors that can be used to produce phenol, ketone, and organic carbonate intermediates include gas fermentation reactors, Fischer-Tropsch reactors, and oxidative carbonylation reactors. In some cases, especially for the reaction pathways used to form phenol compounds, system 801 may include multiple intermediate modules or multiple reactors. In one example, system 801 includes one module for producing simple liquid hydrocarbons, another module for decomposing these hydrocarbons to produce aromatic and other unsaturated carbon-containing compounds, and / or one or more additional intermediate reactors for producing ketones, organic carbonates, and / or phenol derivatives. In various embodiments, these intermediate modules for producing ketones and / or phenols use a combination of carbon monoxide and excess hydrogen produced by carbon dioxide reduction electrolyzer 803.

[0259] FIG. 8B shows an example of a polycarbonate synthesis system 811 having a carbon dioxide reduction electrolyzer 813 configured to receive carbon dioxide and water as reactant inputs and electricity and to drive an electrolysis reaction at an anode and a cathode. The carbon dioxide electrolyzer 813 is configured to output carbon monoxide. As shown, system 811 is configured to deliver the carbon monoxide output by electrolyzer 813 to a carbon monoxide purification module 812.

[0260] In certain embodiments, a carbon monoxide purification unit such as 812 is configured as described above, such as any one of the embodiments described in connection with the Fischer-Tropsch process. See, for example, the CO purification units of FIGS. 19 and 20.

[0261] In certain embodiments, system 811 is configured to provide waste heat from electrolyzer 813 to carbon monoxide purification unit 812 to facilitate the purification of carbon monoxide.

[0262] System 811 is further configured to deliver purified carbon monoxide from the carbon monoxide purification unit 812 to the phosgene production reactor 815. As shown, the phosgene production reactor 815 is configured to receive chlorine gas in addition to the purified carbon monoxide. The phosgene production reactor 815 is configured to produce phosgene, which is provided to the polycarbonate production reactor 819 via appropriate transport components during operation.

[0263] In the illustrated embodiment, the system 811 includes a gas fermentation subsystem 817 configured to receive carbon monoxide and hydrogen gas from the carbon dioxide electrolyzer 813 as reactants. In the illustrated embodiment, the gas fermentation subsystem 817 is configured to react carbon monoxide and hydrogen to produce acetone.

[0264] The gas fermentation subsystem 817 is also configured to produce carbon dioxide as an output. In a particular configuration, the system 811 is configured to directly provide excess carbon dioxide from the output of the subsystem 817 to the electrolyzer 813. In some implementations, the system 811 is configured to directly provide carbon dioxide to the feedstock for the electrolyzer 813.

[0265] As shown, the system 811 is configured to transport acetone from the gas fermentation subsystem 817 to the bisphenol A production unit 814. The bisphenol A production unit 814 is configured to react acetone and phenol to produce bisphenol A. Acetone enters from the gas fermentation subsystem 817 as described above. Phenol may be provided from any of a variety of sources, including those that use carbon monoxide or other outputs from the electrolyzer 813.

[0266] In the illustrated embodiment, bisphenol A from reactor 814 is provided to polycarbonate production reactor 819 via appropriate transfer components during operation. Additionally, phosgene from reactor 815 is provided to reactor 819 via transfer components during operation, and the reactor is configured to react phosgene and bisphenol A to produce polycarbonate as a final output.

[0267] In the illustrated embodiment, system 811 is configured using heat exchangers and / or other heat transfer components to provide heat as needed between various intermediate reactors and subsystems. For example, bisphenol a synthesis reactor 814 and phosgene reactor 815 may be configured to transfer heat between them as needed during various reaction stages.

[0268] FIG. 8C is a diagram showing polycarbonate production system 821. The illustrated system includes a carbon dioxide reduction electrolyzer 823 configured to provide a carbon monoxide output to a carbon monoxide purification unit 822, and the carbon monoxide purification unit is configured to provide purified carbon monoxide to a phosgene production reactor 825. Phosgene production reactor 825 is configured to produce phosgene and output it to a polycarbonate synthesis reactor 829 that produces a final polycarbonate polymer. The components of this phosgene production pathway are generally constructed and may be operated as described in conjunction with other polycarbonate production embodiments described herein.

[0269] In the embodiment shown in FIG. 8C, bisphenol A is provided via a pathway that receives carbon monoxide and hydrogen gas from electrolyzer 823 and converts these input gases to phenol and acetone via a naphtha production reaction.

[0270] In the illustrated embodiment, system 821 is configured to transport a portion of the carbon monoxide produced by electrolyzer 823, along with the hydrogen gas produced by electrolyzer 823, to reactor 827 configured to produce naphtha. Reactor 827 is, in certain embodiments, a Fischer-Tropsch reactor. In other embodiments, reactor 827 is a gas fermentation reactor configured to produce naphtha from carbon monoxide and hydrogen inputs. Regardless of which alternative reactor is used, the output is naphtha. Naphtha is a mixture of various hydrocarbons that can contain, for example, straight-chain, branched, and / or cyclic aliphatic hydrocarbons having from about 5 to 10 carbon atoms. System 821 is further configured to provide any excess carbon dioxide optionally produced by reactor 827 to carbon dioxide electrolyzer 823 for combination with the input carbon dioxide feedstock to the electrolyzer cathode.

[0271] System 821 is configured to deliver naphtha from reactor 827 to naphtha cracking unit 826 configured to operate in a mode to convert the reactant naphtha into various unsaturated hydrocarbons such as toluene, benzene, and propylene. In the illustrated embodiment, the waste from the naphtha cracking reaction performed in reactor 826 includes hydrogen and waste heat. In the illustrated embodiment, system 821 is configured to provide waste heat and hydrogen gas from naphtha cracking unit 826 to Fischer-Tropsch reactor or gas fermentation reactor 827.

[0272] System 821 is configured to provide the unsaturated hydrocarbon output of cracking unit 826 to one or more reactors for converting these unsaturated hydrocarbons into phenol and / or acetone. In the illustrated embodiment, system 821 includes phenol synthesis reactor 828 configured to receive toluene from naphtha cracking reactor 826 and convert the toluene into phenol.

[0273] Various processes may be used to convert toluene to phenol. One of these involves oxidizing toluene to benzoic acid with atmospheric oxygen, which is carried out in the liquid phase at a temperature of about 100 - 150 °C and an absolute pressure of about 3 bar. Cobalt naphthenate is used as a soluble catalyst at a concentration of 0.1 - 0.3%. In a second step, the oxidation of benzoic acid with atmospheric oxygen and steam uses molten benzoic acid as the reactant and solvent at a temperature of about 230 - 240 °C and atmospheric pressure. Copper(II) benzoate is used as the soluble catalyst. A magnesium salt may be added to act as a promoter. In this reaction, copper(II) benzoate decomposes to copper(I) benzoate and benzoyl salicylic acid (2-(benzoyloxy)benzoic acid). Copper(I) benzoate is regenerated to copper(II) benzoate using atmospheric oxygen. Benzoyl salicylic acid is hydrolyzed with steam to form benzoic acid and salicylic acid (2-hydroxybenzoic acid). Salicylic acid is rapidly decarboxylated to phenol and carbon dioxide.

[0274] Additionally, system 821 is configured using components for transporting benzene and propylene from naphtha cracker reactor 826 to a cumene process reactor or subsystem 828' configured to react benzene and propylene to produce phenol and acetone. In some implementations, subsystem 828' is configured to react benzene and propylene via an alkylation reaction in the presence of phosphoric acid and a catalyst to produce cumene, and can then be reacted in the presence of oxygen and sulfuric acid to produce phenol and acetone (Hock rearrangement).

[0275] System 821 further includes a bisphenol A production reactor 824 configured to receive phenol and acetone from phenol production reactor 828 and cumene process reactor / subsystem 828'. Bisphenol A synthesis reactor 824 is configured to produce bisphenol A from phenol and acetone reactants.

[0276] As described above, system 821 further includes a polycarbonate synthesis reactor 829. Reactor 829 is configured to receive bisphenol A from reactor 824 and phosgene from reactor 825 as described above and to produce a polycarbonate output.

[0277] As shown, system 821 is configured to transfer heat as needed between various components such as between cumene process reactor / subsystem 828' and phosgene reactor 825.

[0278] FIG. 8D shows a polycarbonate synthesis system 831 that is similar to the system 821 of FIG. 8C in some implementations but has a parallel path from electrolyzer 823 to bisphenol A synthesis reactor 824 to deliver an acetone input to reactor 824.

[0279] In the embodiment of FIG. 8D, system 831 includes a gas fermentation reactor 837' configured to receive carbon monoxide and hydrogen gas from electrolyzer 833 and produce acetone via a biological fermentation reaction. In the illustrated embodiment, system 831 is configured to transfer acetone from fermentation reactor 837' to bisphenol A synthesis reactor 834.

[0280] System 831 further includes a phosgene path configured to receive carbon monoxide produced by carbon dioxide electrolyzer 833 and output phosgene to polycarbonate production reactor 839. The phosgene path includes, as shown, a carbon monoxide purification unit 832 configured to receive carbon monoxide from electrolyzer 832 and a phosgene production reactor 835 configured to receive purified carbon monoxide from unit 832.

[0281] System 831 further includes a phenol production path configured to receive carbon monoxide and hydrogen gas produced by carbon dioxide electrolyzer 833 and output phenol to bisphenol A production reactor 834.

[0282] In the phenol pathway, system 831 includes a reactor gas fermentation reactor or Fischer-Tropsch reactor 837 configured to produce naphtha from carbon monoxide and hydrogen generated in an electrolytic cell. System 831 is further configured to provide excess carbon dioxide optionally generated by reactor 837 to carbon dioxide electrolytic cell 833.

[0283] System 831 further includes a naphtha cracking unit 836 configured to convert naphtha from reactor 837 into various unsaturated hydrocarbons such as toluene, benzene, and propylene. System 831 is configured to deliver hydrogen and heat generated by the naphtha cracking reaction in reactor 836 to a Fischer-Tropsch reactor or gas fermentation reactor 837.

[0284] System 831 is configured to provide the unsaturated hydrocarbon output of cracking unit 836 to one or more reactors configured to convert these unsaturated hydrocarbons into phenol and / or acetone. In the illustrated embodiment, system 831 includes a phenol synthesis reactor 838 configured to receive toluene from naphtha cracking reactor 836 and convert the toluene into phenol. Further, system 831 includes a cumene process reactor or subsystem 838' configured to react benzene and propylene from reactor 836 to produce phenol and acetone.

[0285] Phenol and acetone from reactor 838', acetone from gas fermentation reactor 837', and phenol from phenol synthesis reactor 838 are provided to bisphenol A synthesis reactor 834. System 831 is configured to provide bisphenol A from reactor 834 and phosgene from reactor 835 to polycarbonate production reactor 839 and is configured to act on these inputs to produce a polycarbonate polymer.

[0286] FIG. 8E is a diagram showing a polycarbonate production system 841 including a polycarbonate synthesis reactor 849 configured to receive phosgene from carbon monoxide generated in an electrolytic cell and receive bisphenol A from a reactor 844 that receives acetone generated by fermentation. Further, the system 841 includes a chlorine generation path for providing electrolytically generated chlorine to a phosgene production reactor 845.

[0287] In the illustrated embodiment, the system 841 includes a carbon dioxide reduction electrolytic cell 843 that operates and may be configured in a manner similar to the electrolytic cells described in other polycarbonate production systems herein. As shown, the system 841 is configured to transport carbon monoxide generated by the electrolytic cell 843 directly to a carbon monoxide purification unit 842. The system 841 is also configured to transport purified carbon monoxide from the carbon monoxide purification unit 842 to a phosgene production reactor 845. The phosgene production path may operate in a manner similar to other polycarbonate production systems described herein and may be configured to use components similar to those described in other polycarbonate production systems herein. However, in the illustrated embodiment, the chlorine used in the phosgene production reactor 845 is generated electrolytically in conjunction with the operation of the carbon dioxide reduction electrolytic cell 842. In certain embodiments, the chlorine is generated by a chlor-alkali cell that uses a chloride salt (e.g., NaCl) as a source of chloride ions for electrolytic oxidation to produce chlorine gas.

[0288] In the illustrated embodiment, system 841 is further configured to provide carbon monoxide and hydrogen gas from electrolyzer 843 to a gas fermentation reactor 847 configured to convert carbon monoxide and hydrogen gas into acetone via a biological fermentation reaction. System 841 is further configured to convey acetone from fermentation reactor 847 to bisphenol A production reactor 844. As shown, bisphenol A production reactor 844 is configured to receive phenol in addition to acetone as inputs and react them to produce bisphenol A. As shown, system 841 is configured to deliver bisphenol A from reactor 844 to polycarbonate production reactor 849. Further, system 841 is configured to transfer heat between bisphenol A production reactor 844 and phosgene production reactor 845 as needed during the polycarbonate production process.

[0289] FIG. 8F is a diagram showing a polycarbonate production system 851 including an electrolysis subsystem that includes a carbon dioxide reduction electrolyzer 853 and a chlor-alkali system 853'. Chlor-alkali system 853' is configured to receive water and sodium chloride as inputs and produce chlorine gas and hydrogen gas as outputs. System 851 is configured to deliver chlorine produced by chlor-alkali system 53' to phosgene production reactor 855. System 851 is also configured to deliver hydrogen gas optionally produced by chlor-alkali system 853' to gas fermentation reactor 857. In certain embodiments, system 851 is configured to supply oxygen to the cathode of chlor-alkali system 853' including an oxygen depolarized cathode from electrolyzer 853. In certain embodiments, system 851 includes a water electrolyzer configured to produce oxygen that can be delivered to the oxygen depolarized cathode of the chlor-alkali cell. System 851 may also be configured to provide hydrogen from the water electrolyzer to a gas fermentation reactor or other reactor used to produce diol.

[0290] System 851 is further configured to deliver carbon monoxide and hydrogen gas from the carbon dioxide reduction electrolyzer 853 to the gas fermentation reactor 857. Thus, the gas fermentation reactor 857 is configured to receive hydrogen from both the carbon dioxide reduction electrolyzer 853 and the chlor-alkali system 853'. The gas fermentation reactor 857 is configured to perform biological fermentation on the carbon monoxide and hydrogen gas input and produce acetone as an output. The gas fermentation reactor 857 is also configured to produce carbon dioxide as a byproduct. In the illustrated embodiment, system 851 is configured to deliver excess carbon dioxide produced by reactor 857 to electrolyzer 853.

[0291] System 851 is further configured to transport acetone produced by the gas fermentation reactor 857 to the bisphenol A production reactor 854. The bisphenol A production reactor 854 is also configured to include an input for receiving phenol. The bisphenol A production reactor 854 is configured to react acetone and phenol to produce bisphenol A.

[0292] Another path of system 851 is a phosgene production path that includes a carbon monoxide purification unit 852 configured to receive and purify carbon monoxide produced by the carbon dioxide reduction electrolyzer 853. System 851 is further configured to provide purified carbon monoxide from the purification unit 852 to the phosgene production reactor 855. As described above, the phosgene production reactor 855 is also configured to receive chlorine from the chlor-alkali system 855'.

[0293] As shown, the polycarbonate production system 851 further includes a polycarbonate production reactor 859, as well as components for transporting bisphenol A from the bisphenol A production reactor 854 and transporting phosgene from the phosgene production reactor 855 to the polycarbonate production reactor 859. The reactor 859 is configured to react bisphenol A and phosgene to produce a polycarbonate polymer.

[0294] FIG. 8G is a diagram showing a polycarbonate production system 861 that uses three separate pathways from a carbon dioxide reduction electrolyzer subsystem. The first carbon dioxide reduction electrolyzer 863 is configured to (a) produce carbon monoxide and hydrogen for the naphtha pathway and (b) produce carbon monoxide for the phosgene pathway. A separate carbon dioxide reduction electrolyzer 863' is configured to produce acetone for the acetone pathway.

[0295] The phosgene production pathway may be similar to that described in other systems for producing polycarbonates. This includes a carbon monoxide purification unit 862 and a phosgene production reactor 865. The carbon monoxide purification unit 862 is configured to receive carbon monoxide and waste heat from the electrolyzer 863. It is also configured to provide purified carbon monoxide to a phosgene production reactor 865 that has inputs for receiving chlorine gas as well as phosgene.

[0296] The acetone pathway includes a separate carbon dioxide reduction electrolyzer 863' that is designed and / or operated in a different way than the electrolyzer 863, as described above. As shown, the system 861 is configured to deliver acetone directly from the electrolyzer 863' to a bisphenol A synthesis reactor 864.

[0297] Similar to the polycarbonate production system 831 shown in FIG. 8D, the naphtha pathway includes a Fischer-Tropsch reactor or gas fermentation reactor 867, a naphtha cracking subsystem 866, a phenol production reactor 868, and a cumene process reactor 868'. The Fischer-Tropsch reactor or gas fermentation reactor is configured to receive carbon monoxide and hydrogen gas from the carbon dioxide electrolyzer 863 and output naphtha. The Fischer-Tropsch reactor or gas fermentation reactor 867 is also configured to receive hydrogen gas from the naphtha cracking reactor 866 and return excess carbon dioxide to the electrolyzer 863.

[0298] The naphtha cracking subsystem 866 is configured to produce at least propylene, benzene, and toluene. The system 861 is configured with a conveyance component for delivering benzene and propylene from the naphtha cracking subsystem 866 to a cumene process reactor 868 configured to generate phenol and acetone as outputs. The system 861 is also configured to convey toluene from the naphtha cracking subsystem 866 to a phenol production reactor 868 configured to produce phenol. The system 861 is further configured to deliver acetone and phenol from the cumene process reactor 868' to a bisphenol A synthesis reactor 864 together with the phenol produced by the phenol production reactor 868. As described above, the system 861 is also configured to deliver acetone from the electrolytic cell 863' to the bisphenol A synthesis reactor 864.

[0299] The system 861 is further configured to deliver the bisphenol A produced by the reactor 868 to a polycarbonate synthesis reactor 869 together with the phosgene produced by the reactor 865.

[0300] FIG. 8H shows another implementation of a polycarbonate polymer production system. This system is designated as 871. It includes a single carbon dioxide reduction electrolytic cell 873 and components for conveying carbon monoxide and hydrogen gas from the electrolytic cell 873 to a gas fermentation reactor and conversion reactor or subsystem 877 configured to directly produce phenol and excess carbon dioxide. The system 871 is configured to return the excess carbon dioxide produced by the reactor or subsystem 877 to the inlet on the cathode side of the electrolytic cell 873.

[0301] Reactor or subsystem 877 may be configured to produce phenol from a gas input, either alone or in combination with sugar and / or biomass. In some implementations, the gas fermentation reactor produces mevalonic acid or mevalonolactone, which can be converted to phenol by heating a silica catalyst in a precursor. In certain embodiments, the microorganism used to produce mevalonic acid is a naturally occurring microorganism such as Escherichia coli that has been modified to express the MVL pathway.

[0302] System 871 is further configured to deliver phenol from reactor 877 to bisphenol A synthesis reactor 874, which is also configured to receive acetone and phenol as inputs and produce bisphenol A as an output.

[0303] System 871 is also configured to transport carbon monoxide produced by electrolyzer 873 to carbon monoxide purification unit 872 and transport purified carbon monoxide from unit 872 to phosgene production reactor 875, which is configured to receive purified carbon monoxide along with chlorine and produce phosgene.

[0304] System 871 is further configured to transport phosgene from reactor 875 and bisphenol A from reactor 874 to polycarbonate synthesis reactor 879, which reacts bisphenol A and phosgene to produce and output a polycarbonate polymer.

[0305] In certain embodiments, polycarbonate synthesis is carried out without phosgene, but nevertheless using carbon monoxide generated from a carbon dioxide electrolyzer. The polycarbonate synthesis system can be configured for various non-phosgene routes to polycarbonate. In some phosgene-free routes, the polymerization relies on the transesterification of DPC (diphenyl carbonate) and bisphenol A. In certain embodiments, the non-phosgene system is configured to produce an intermediate dialkyl carbonate, such as dimethyl carbonate (DMC), as a source of carbonate functional groups. These systems may be configured to react phenol with dimethyl carbonate to produce, for example, phenyl methyl carbonate. Various non-phosgene routes use methods for producing dialkyl carbonates. In certain embodiments, these are produced using carbon monoxide from a carbon dioxide electrolyzer. As an example, DMC may be produced using oxidative carbonylation. CO + 1 / 2O2 + 2CH3OH → (CH3O)2CO + H2O Formate

[0306] Alkali metal formates have many uses, including as enzyme stabilizers in liquid detergents. The enzyme can be lipase, amylase, protease, etc. Other formates, such as alkaline earth metal formates, also have many uses. In certain embodiments, the formate production system uses a carbon dioxide reduction electrolyzer to convert carbon dioxide to carbon monoxide, which is then processed to produce an alkali metal formate. In various embodiments, the metal formate is produced by contacting a metal hydroxide with carbon monoxide. The contact can occur in a liquid (e.g., aqueous) medium or a solid medium.

[0307] FIG. 9 is a diagram showing an exemplary formate production system 901 comprising a carbon dioxide reduction electrolyzer 903, a formate production reactor 905, and various downstream formate recovery units. The electrolyzer 903 is configured to receive oxygen and carbon dioxide as reactants and receive electricity to promote the reduction of carbon dioxide to produce carbon monoxide. The system 901 is configured to transport carbon monoxide from the electrolyzer 903 to the formate production reactor 905, where the carbon monoxide reacts with a hydroxide (e.g., sodium hydroxide, potassium hydroxide, cesium hydroxide, or calcium hydroxide) to produce a dissolved metal formate. The reactor 905 is configured to receive not only carbon monoxide from the electrolyzer 903, but also a metal hydroxide, a solvent, and a catalyst. The reactor 905 may be a stirred tank reactor.

[0308] In certain embodiments, a carbon dioxide electrolyzer located upstream from the metal formate production reactor is configured to operate in (a) the operating parameter regime of high reduction product to hydrogen production stream described herein, and / or (b) the operating parameter regime of high reduction product to CO2 ratio described herein.

[0309] System 901 is configured to transport a formate-containing solution from reactor 905 to degassing unit 907, which removes gas from the formate solution during operation. Such gas includes unreacted carbon monoxide. System 901 is further configured to transport the degassed formate solution from unit 907 to evaporator 909 configured to at least partially evaporate a solvent from the formate solution to produce a slurry or other liquid-solid mixture containing precipitated metal formate. System 901 further includes a filtration unit 911 configured to receive and filter the output of evaporator 909. The output of filtration unit 911 includes concentrated solid metal formate. System 901 further includes a solvent wash unit 913 configured to wash the solid formate-containing output of unit 911 by contacting the formate material with a solvent. System 901 is further configured to transport the filtered and washed solid formate from unit 913 to a dryer 915 configured to dry the solid formate to produce the final form of solid metal formate. Dryer 915 is configured to receive a drying gas such as nitrogen or carbon dioxide along with heat. In certain embodiments, dryer 915 is configured to receive waste heat from electrolyzer 903. In certain embodiments, dryer 915 is configured to receive carbon dioxide from electrolyzer 903 or from an inlet stream to electrolyzer 903. In some implementations, externally provided dried carbon dioxide is sent from dryer 915 to the input stream of electrolyzer 903.

[0310] In some implementations, the formate produced by system 901 is an alkali metal formate such as sodium, potassium, or cesium formate, or an alkaline earth metal formate such as calcium formate or barium formate. Optionally, system 901 is configured to produce formic acid from the metal formate by using a reactor configured to contact the metal formate with an acid such as hydrochloric acid.

[0311] In some embodiments, the metal formate is produced by contacting carbon monoxide produced by an electrolytic carbon dioxide cell with a solid or slurry of metal hydroxide. For example, sodium formate can be produced by contacting solid sodium hydroxide with a stream of carbon monoxide. The reaction can be represented as NaOH(s) + CO(g) -> NaCOOH(s). The solid hydroxide can be provided in various forms such as a powder. Optionally, in some cases, the particle size is reduced by grinding, pulverizing, or other methods during the reaction with carbon monoxide to increase the surface area of the hydroxide available for the reaction. For example, the solid hydroxide may be ground in a ball mill autoclave during contact with carbon monoxide. Optionally, in some cases, the solid hydroxide is contacted with carbon monoxide at a temperature of at least about 200 °C (e.g., about 230 to 300 °C) and / or at a pressure of at least about 2 bar (e.g., about 5 to 10 bar) during the reaction in the autoclave. In some embodiments, the formate production reaction in the autoclave has a residence time of at least about 15 - 60 minutes or about 20 - 40 minutes.

[0312] In some metal formate syntheses, carbon monoxide is provided to a reactor (e.g., an autoclave containing a solid metal hydroxide) at a concentration of at least about 0.5 mole fraction, or at least about 0.8 mole fraction, or at least about 0.9 mole fraction. Ethylene glycol

[0313] Figures 10A and 10B illustrate a process for preparing ethylene glycol (monoethylene glycol or MEG). As shown in Figure 10A, the MEG production system 1001 includes a carbon dioxide electrolyzer 1003, an ethylene oxide production reactor 1005, and an MEG production reactor 1007. The electrolyzer 1003 is configured to produce ethylene. The system 1001 is configured to deliver ethylene from the electrolyzer 1003 to the ethylene oxide production reactor 1005. Optionally, the system 1001 is further configured to deliver oxygen from the electrolyzer 1003 to the reactor 1005. Regardless of the source of oxygen, the reactor 1005 is configured to react ethylene and oxygen to produce ethylene oxide.

[0314] In certain embodiments, the reactor 1005 is a direct ethylene oxide reactor designed or configured to produce ethylene oxide directly from ethylene and oxygen. In this approach, ethylene and compressed oxygen may be fed to a multitubular catalytic reactor (an example of reactor 1005). During operation of such a reactor, the mixture is sent over a silver oxide catalyst supported on a porous carrier at about 200 - 300 °C and about 10 - 30 bar. The reaction is exothermic and the heat removed can be used elsewhere in the system. The system 1001 may be configured to cool the gases from the reactor 1005 and pass them through a scrubber where ethylene oxide is absorbed as a dilute aqueous solution.

[0315] The system 1001 may be configured to deliver ethylene oxide from the reactor 1005 to the ethylene glycol production reactor 1007. The reactor 1007 may be configured to react ethylene oxide and water to produce ethylene glycol. The reaction may be catalyzed by an acid or a base, or may be carried out at neutral pH and elevated temperature. In certain embodiments, the system 1001 is configured to provide heat from the electrolyzer 1003 and / or the ethylene oxide production reactor 1005 to the MEG reactor 1007.

[0316] As shown in FIG. 10B, the MEG generation system 1011 includes a carbon dioxide electrolyzer 1013 and a MEG generation reactor 1017. The electrolyzer 1013 is designed or configured to produce carbon monoxide and hydrogen. The system 1011 is configured to deliver these outputs to the reactor 1017 along with oxygen (optionally from the electrolyzer 1013), where the reactants react to produce ethylene glycol. The reactor 1017 may be configured to produce ethylene glycol from these reactants via a two-step process that produces dimethyl oxalate as an intermediate from a reaction pathway that includes methanol, dinitrogen trioxide, and carbon monoxide. The production of dimethyl oxalate may use a palladium catalyst. The reactor 1017 may be configured to perform a second step by reacting dimethyl oxalate with hydrogen gas using a copper catalyst to produce ethylene glycol. In this process, only carbon monoxide, hydrogen, and oxygen are consumed. Hydrogen for the reaction may come in from any suitable source. A common source of hydrogen is shown as the hydrogen generator 1019 of the system 1011. In certain embodiments, the hydrogen generator 1019 is a water electrolyzer. In certain embodiments, the hydrogen generator 1019 is a reactor configured to perform an aqueous shift reaction. In certain embodiments, hydrogen is produced from fossil fuels and the carbon dioxide product is optionally recycled to the electrolyzer 1013. The system 1011 may be configured to provide excess oxygen from the electrolyzer 1013 to a combustion reactor.

[0317] In certain embodiments, a carbon dioxide electrolyzer disposed upstream from the MEG generation reactor is configured to operate in (a) the operating parameter regime of the hydrogen-rich product stream described herein, and / or (b) the operating parameter regime of the high reduction product to CO2 ratio described herein. Polyethylene terephthalate

[0318] FIG. 11 is a schematic diagram of a system 1101 that can be used to produce the polymer polyethylene terephthalate. One process or group of processes uses an electrolytic carbon dioxide cell 1103 configured to produce ethylene. Another process or group of processes uses an electrolytic carbon dioxide cell 1113 configured to produce carbon monoxide and hydrogen.

[0319] System 1101 is configured to implement a PET production pathway that includes electrolytic cell 1103. As shown, system 1101 further includes an ethylene oxide production reactor 1105 and an ethylene glycol production reactor 1107 that can be configured and arranged like system 1001 of FIG. 10A. System 1101 is configured to deliver ethylene glycol to a PET production reactor 1109 configured to react ethylene glycol with phthalic acid to produce a PET polymer.

[0320] The type of system 1101 for producing both ethylene glycol and phthalic acid includes an electrolytic cell 1113 configured to produce carbon monoxide and hydrogen. The ethylene glycol production pathway is configured to react these products, optionally, with the addition of excess hydrogen from a source 1119 within a MEG production reactor 1117. System 1101 is configured to provide MEG from reactor 1117 to PET production reactor 1109. The type of system 1101 that uses electrolytic cell 1113 optionally does not include components for a MEG pathway that uses electrolytic cell 1103, ethylene oxide production reactor 1105, and MEG synthesis reactor 1107.

[0321] The type of system 1101 using electrolyzer 1113 may also include a reactor for producing terephthalic acid from carbon monoxide and hydrogen generated by electrolyzer 1113. The reactor may produce naphtha and p-xylene as intermediates. In the illustrated embodiment, the Fischer-Tropsch reactor or gas fermentation reactor 1121 is configured to produce naphtha from carbon monoxide and hydrogen output by electrolyzer 1113. Reactor 1121 may be designed or configured as described elsewhere in this specification. The naphtha cracking reactor 1123 is configured to crack naphtha to produce p-xylene. In certain embodiments, system 1101 is configured to supply excess heat and / or hydrogen generated by cracker 1123 to reactor 1121. The PTA reactor 1125 is configured to convert p-xylene to terephthalic acid.

[0322] In certain embodiments, the carbon dioxide electrolyzer disposed upstream from the Fischer-Tropsch reactor is configured to operate in (a) the operating parameter regime of the hydrogen-rich product stream described herein, and / or (b) the operating parameter regime of the high reduction product to CO2 ratio described herein.

[0323] In some types of system 1101, the carbon dioxide reduction electrolyzer is configured to directly produce ethylene glycol (MEG) by electrolysis. In such types, the electrolyzer replaces or supplements another MEG production route, such as one that uses reactor 1107 or reaction 1117. Some types of system 1101 use a reactor configured to directly convert carbon monoxide from reactor 1113 to p-xylene together with hydrogen. This type of system 1101 is configured to transport such p-xylene to PTA reactor 1125. Acetic acid

[0324] FIG. 12 schematically shows a system 1201 for producing acetic acid from carbon monoxide and hydrogen produced by a carbon dioxide electrolyzer 1203. The system 1201 includes a methanol production reactor 1205 configured to react carbon monoxide and hydrogen to produce methanol. The reactor 1205 may be configured like a conventional methanol synthesis reactor using synthesis gas.

[0325] In certain embodiments, the carbon dioxide electrolyzer disposed upstream from the methanol production reactor is configured to operate in (a) the operating parameter regime of the hydrogen-rich product stream described herein, and / or (b) the operating parameter regime of the high reduction product to CO2 ratio described herein.

[0326] The system 1201 also includes an acetic acid production reactor 1207 configured to react methanol and purified carbon monoxide to produce acetic acid. The reactor 1207 may be configured to perform methanol carbonylation using, for example, a metal carbonyl catalyst. In certain embodiments, the system 1201 includes a carbon monoxide purification unit 1209 configured to produce purified carbon monoxide. The carbon dioxide purification unit may be designed by the methods described elsewhere herein (e.g., by the methods of the units of FIGS. 19 and 20). Isocyanate

[0327] FIG. 13 schematically shows a system 1301 configured to produce diisocyanate from electrolytically produced carbon monoxide.

[0328] System 1301 may be configured to transport carbon monoxide produced by carbon dioxide electrolyzer 1313 to carbon monoxide purification unit 1317. System 1301 may also be configured to transport purified carbon monoxide from unit 1317 to phosgene production reactor 1319 configured to react the purified carbon monoxide with chlorine to produce phosgene. In certain embodiments, phosgene production reactor 1319 is designed or configured to operate in a manner similar to other phosgene production reactors described herein, such as in relation to a polycarbonate production system. Examples of carbon monoxide purification units are presented in FIGS. 19 and 20 and the related description.

[0329] In certain embodiments, the carbon dioxide electrolyzer disposed upstream from the phosgene production reactor is configured to operate in (a) the operating parameter regime of the high reduction product to hydrogen production stream described herein, and / or (b) the operating parameter regime of the high reduction product to CO2 ratio described herein.

[0330] In certain embodiments, system 1301 includes an electrolytic cell, such as a chlor-alkali cell, for producing chlorine. System 1301 may be configured to provide chlorine to phosgene production reactor 1319. In some implementations, system 1301 is configured to supply the oxygen byproduct of electrolyzer 1313 to the cathode of a chlor-alkali cell including an oxygen depolarized cathode. Oxygen from electrolyzer 1313 may be used in place of oxygen from other sources such as air separation. In some implementations, system 1301 includes a water electrolyzer, and system 1301 is configured to supply oxygen produced by the water electrolyzer to the oxygen depolarized cathode of the chlor-alkali cell. System 1301 may also be configured to provide hydrogen from the water electrolyzer to an amine production reactor.

[0331] In the illustrated embodiment, system 1301 is configured to transport phosgene from phosgene generator reactor 1319 to isocyanate generator reactor 1321 configured to react phosgene with an amine to produce a polyisocyanate, e.g., a diisocyanate such as toluene diisocyanate (TDI) or methylene diisocyanate (MDI) depending on the structure of the amine being supplied. In some implementations, reactor 1321 is configured to react phosgene with a free amine in an inert organic solvent at low temperature. The resulting mixture of carbamoyl chloride and amine hydrochloride is then reacted at a higher temperature to produce the desired polyisocyanate.

[0332] In certain embodiments, the amine reactant is produced by a reactor or reaction that uses one or more carbon oxide reduction products produced by the carbon oxide electrolyzer described herein, e.g., via a Fischer-Tropsch reaction and a cracking reaction. In certain embodiments, the amine reactant is produced by a bioreactor such as a gas fermentation reactor. In certain embodiments, system 1301 is configured to provide electro-generated hydrogen (optionally from electrolyzer 1313 or a water electrolyzer) to a gas fermentation reactor configured to produce an amine product or an intermediate used in amine production. In some implementations, the hydrogen for amine production is provided by a separate source.

[0333] Regardless of the amine source, reactor 1321 may be configured to react a polyamine with phosgene by a phosgenation reaction to produce a polyisocyanate such as a diisocyanate. In certain embodiments, the diisocyanate is 2,4-toluene diisocyanate and / or 2,6-toluene diisocyanate. In certain particular embodiments, the diisocyanate is 4,4'-diphenylmethane diisocyanate.

[0334] In various embodiments, substantially pure carbon monoxide is used to produce phosgene, which is then reacted with various amines to produce isocyanates. Some amines may be produced using hydrogen, thus providing a use for byproduct hydrogen from a carbon dioxide electrolyzer or a water electrolyzer located in the same location.

[0335] As an example, in the case of a system configured to produce toluene diisocyanate (TDI), the carbon dioxide electrolyzer may be configured or operated to produce a hydrogen-rich output stream. A high-hydrogen-containing stream (e.g., an H2:CO ratio of about 1:1) may be used for the production of the precursor toluene diamine (TDA) and phosgene. TDA may be produced by the hydrogenation of dinitrotoluene. For example, hydrogen can be used to produce nitric acid (described elsewhere), which is used for the nitration of toluene to produce dinitrotoluene.

[0336] In certain embodiments, a system comprising a carbon dioxide electrolyzer is configured to produce methylene diisocyanate (MDI) from aniline and phosgene. Aniline may be produced by the hydrogenation of nitrobenzene. In some embodiments, a system for producing MDI via aniline is configured to produce a feed gas with a relatively high ratio of hydrogen to carbon monoxide (e.g., hydrogen:CO around about 3:1). In some embodiments, the system is configured to produce nitric acid (described elsewhere) that is used for the nitration of benzene to produce nitrobenzene. The system may be configured to use a separate gas stream containing a relatively low concentration of hydrogen (e.g., a ratio of hydrogen to CO around about 1:1), which can be used to produce formaldehyde, which is then reacted with aniline to produce a diamine, which is subsequently phosgenated to produce MDI.

[0337] In certain embodiments, a system comprising a carbon dioxide electrolyzer is configured to produce hexamethylene diisocyanate (HDI) using a hydrogen-enriched gas stream (e.g., a gas stream having an H2:CO ratio of about 4:1). The system may be configured to hydrogenate adiponitrile to produce hexamethylenediamine, which is then phosgenated to produce HDI. Polyurethane

[0338] FIG. 14 provides a schematic diagram of a system 1401 that can be used to produce polymeric polyurethane. A group of reactors uses a carbon dioxide electrolyzer 1403 configured to produce ethylene. Another group of reactors uses a carbon dioxide electrolyzer 1413 configured to produce carbon monoxide and hydrogen. In an alternative embodiment, only one of these two groups of reactors is used, and an alternative source is used to provide the intermediate chemical that would otherwise be produced by the other group of reactors.

[0339] In some implementations, system 1401 is configured to transport ethylene, and optionally oxygen, from electrolyzer 1403 to ethylene oxide production reactor 1405. System 1401 is also configured to provide ethylene oxide from reactor 1405 to ethylene glycol production reactor 1407 and / or polyethylene glycol production reactor 1409. In certain embodiments, ethylene oxide production reactor 1405 and / or ethylene glycol production reactor 1407 are designed or configured to operate in a manner similar to reactors 1005 and 1007 of system 1001.

[0340] The polyethylene glycol production reactor 1409 may be configured to produce polyethylene glycol from the interaction of ethylene oxide with water, ethylene glycol, and / or ethylene glycol oligomers. The length and polydispersity of the PEG chains of the product are affected by the selection and ratio of the reactants. The system 1401 may be configured to transport the heat generated by the exothermic PEG production reaction from the reactor 1409 to a carbon monoxide purification process (e.g., for phosgene production) or to another process that requires energy. Examples of carbon monoxide purification units are presented in FIGS. 19 and 20 and the related description.

[0341] In some implementations, the system 1401 is configured to transport carbon monoxide, and optionally hydrogen, from the electrolyzer 1413 to a reactor or group of reactors 1415 configured to produce one or more polyols (e.g., polyethylene glycol). The reactor or group of reactors 1415 may be a bioreactor configured to produce polyols by a gas fermentation reaction. In some embodiments, the reactor 1415 is configured to produce polyols using an algae-based reaction. In certain embodiments, the reactor or group of reactors 1415 includes a Fischer-Tropsch reactor and / or a naphtha cracking reactor used to produce hydrocarbons that can be converted to polyols.

[0342] In certain embodiments, a carbon dioxide electrolyzer located upstream from the naphtha production and cracking subsystem is configured to operate in (a) the operating parameter regime of the hydrogen-rich product stream described herein, and / or (b) the operating parameter regime of the high reduction product-to-CO2 ratio described herein.

[0343] System 1401 may be configured to utilize carbon monoxide and optionally hydrogen produced by electrolyzer 1413 to produce diisocyanates such as MDI or TDI. In certain embodiments, system 1401 is configured to perform diisocyanate production using system 1301 shown in FIG. 13 as a subsystem. Such a subsystem can include a carbon monoxide purification unit and a phosgene production reactor. Regardless of how diisocyanate precursors such as phosgene and free amines are produced, system 1401 is configured to react them in diisocyanate production reactor 1421. Different types of diisocyanates may be used depending on the polyurethane produced. Examples include 2,4-toluene diisocyanate and / or 2,6-toluene diisocyanate. In certain embodiments, the diisocyanate is 4,4'-diphenylmethane diisocyanate.

[0344] System 1401 may optionally be configured to provide excess carbon monoxide in the outlet stream of ethylene production electrolyzer 1403 to the phosgene / isocyanate production pathway.

[0345] System 1401 includes a polyurethane production reactor 1411 configured to receive a polyol and a diisocyanate and react them to produce a polyurethane polymer. In certain implementations, the polyol is produced, for example, by reactor or reactor group 1415 and / or polyethylene glycol production reactor 1409. In some implementations, system 1401 is configured to transport the polyol from one or both of these reactors and / or transport the diisocyanate from reactor 1421 to polyurethane production reactor 1411. In certain embodiments, reactor 1411 is designed or configured to react upon contact with the vapors of the polyol and diisocyanate. In certain embodiments, the polyol stream includes a catalyst (e.g., and acidic or basic amine), a surfactant, and / or a blowing agent. Oxalic acid

[0346] In certain embodiments, oxalic acid is produced from carbon monoxide generated by a carbon dioxide electrolyzer. Various pathways may be used to produce oxalic acid from carbon monoxide. Examples of systems incorporating these pathways are shown in FIGS. 15-18.

[0347] In certain embodiments, the carbon dioxide electrolyzer and associated oxalic acid production unit are located at or near a plant for producing cement. Carbon dioxide produced by the cement plant may be used as a feedstock for the carbon dioxide electrolyzer. Oxalic acid produced by the system may be used for cement hardening. In certain embodiments, the oxalic acid used in the cement produces calcium oxalate with very low solubility. Cement produced with oxalic acid can withstand degradation due to contact with acids during use (e.g., after installation or construction).

[0348] FIG. 15 shows a system 1501 comprising a carbon dioxide electrolyzer 1503 configured to produce carbon monoxide and hydrogen. System 1501 is configured to transport a portion of the carbon monoxide and hydrogen produced by electrolyzer 1503, optionally with some oxygen, to an alcohol production reactor 1505. Reactor 1505 is configured to produce methanol or butanol in certain embodiments.

[0349] In some embodiments, reactor 1505 is a bioreactor that uses organisms having a metabolic pathway for converting carbon monoxide to an alcohol such as butanol. Examples of such organisms include autotrophic acetylene such as Clostridium Carboxidivorans and Butyribacterium methylotrophicum. In some embodiments, reactor 1505 is configured to produce methanol by catalytic reaction of carbon monoxide and hydrogen at high temperature and high pressure. In some embodiments, the catalyst is a mixture of copper oxide and zinc oxide supported on alumina. In some embodiments, system 1501 is configured to provide additional hydrogen beyond that produced by electrolyzer 1503 to alcohol reactor 1505.

[0350] In certain embodiments, a carbon dioxide electrolyzer disposed upstream from the alcohol production reactor is configured to operate in (a) the operating parameter regime of the hydrogen-rich product stream described herein, and / or (b) the operating parameter regime of the ratio of highly reduced product to CO2 described herein.

[0351] System 1501 is further configured to transport the alcohol produced by reactor 1505 to oxalic acid production reactor 1507. In certain embodiments, reactor 1507 is configured to react the alcohol with carbon monoxide and oxygen to produce a diester oxalate, which is then hydrolyzed to produce free oxalic acid. 4ROH + 4CO + O2 → 2(CO2R)2 + 2H2O

[0352] The carbon monoxide provided to reactor 1507 may be provided directly from electrolyzer 1503. In some embodiments, the carbon monoxide from electrolyzer 1503 is purified before being delivered to reactor 1507.

[0353] Reactor 1507 may produce oxalic acid in an impure form. Thus, system 1501 may be further configured to provide the oxalic acid product to a separator 1509, which may be configured to purify the oxalic acid and return unreacted alcohol to reactor 1507. In certain embodiments, separator 1509 is configured to perform azeotropic distillation on the oxalic acid product from reactor 1507.

[0354] FIG. 16 is a diagram showing a system 1601 configured to produce oxalic acid from carbon monoxide generated by an electrolyzer 1603. As shown, system 1601 is configured to transport carbon monoxide and hydrogen generated by electrolyzer 1603 to an ethylene glycol production reactor 1605. Reactor 1605 may be designed or configured to produce ethylene glycol in a manner similar to that described above with reference to system 1011 of FIG. 10A (see reactor 1017).

[0355] In certain embodiments, a carbon dioxide electrolyzer located upstream from the MEG production reactor is configured to operate in (a) the operating parameter regime of the hydrogen-rich product stream described herein and / or (b) the operating parameter regime of the high reduction product to CO2 ratio described herein.

[0356] System 1601 is configured to transport the ethylene glycol produced by reactor 1605 to an oxalic acid production reactor 1607 configured to oxidize the ethylene glycol to produce oxalic acid. In certain embodiments, reactor 1607 is configured to utilize an oxidizing agent such as nitric acid and / or air to produce oxalic acid from ethylene glycol. In certain embodiments, reactor 1607 is configured to react an alcohol (MEG) in the presence of an oxidizing agent such as air or nitric acid using a catalyst such as vanadium pentoxide to produce oxalic acid.

[0357] In certain embodiments, system 1601 includes a reactor 1609 for producing nitric acid. In certain embodiments, reactor 1609 is configured to carry out the Ostwald process. In certain embodiments, system 1601 additionally includes a reactor for carrying out the Haber process to produce ammonia. System 1601 may be configured to provide ammonia to the Ostwald reactor 1609. In some implementations, a Haber reactor and an Ostwald reactor 1609 are provided as a subsystem that receives hydrogen and nitrogen as reactants and produces nitric acid as a product. System 1601 may be configured to direct the hydrogen produced by electrolyzer 1603 to a subsystem for producing nitric acid (e.g., a subsystem that first produces ammonia from hydrogen and nitrogen). In some embodiments, the nitric acid is supplied from an external source. It should be understood that in other embodiments that require nitric acid, such as other systems for producing oxalic acid, the nitric acid can be produced from a Haber process subsystem that receives hydrogen from an oxidative carbon electrolyzer.

[0358] In the illustrated embodiment, reactor 1607 receives nitric acid from reactor 1609 and produces oxalic acid. In certain embodiments, reactor 1609 produces relatively impure oxalic acid, such as oxalic acid containing a certain amount of nitric acid. In the illustrated embodiment, system 1601 is configured to deliver the impure oxalic acid to a crystallization and separator unit 1611 configured to purify the oxalic acid and return the nitric acid to reactor 1607.

[0359] FIG. 17A shows a system 1701 comprising a carbon dioxide electrolyzer 1703, a metal formate reactor 1705, and an oxalic acid reactor 1707. The system 1701 is configured to deliver carbon monoxide produced by the electrolyzer 1703 to the formate reactor 1705. During operation, the metal hydroxide is provided to the reactor 1705 that produces metal formate, along with heat (optionally, waste heat from the electrolyzer 1703). The production, purification, and / or extraction of formate can proceed as described above with reference to FIG. 9. In some implementations, the formate is an alkali metal formate such as sodium, potassium, or cesium formate, or an alkaline earth metal formate such as calcium formate or barium formate. In a particular embodiment, the formate is sodium formate. In certain embodiments, the formate is potassium formate.

[0360] Regardless of how the metal formate is produced and optionally extracted, the system 1701 is configured to transport the formate to the oxalic acid reactor 1707. The reactor 1707 may be configured to convert the formate to oxalate via a thermal decomposition reaction. The reactor 1707 may also be configured to convert the metal oxalate to oxalic acid upon contact with an acid. In a particular embodiment, the oxalic acid forming reactor 1707 is configured to receive an acid such as hydrochloric acid to facilitate the reaction, for example. In certain embodiments, the halide byproduct (e.g., NaCl) of the oxalate production reaction (within the reactor 1707) is provided to a chlor-alkali electrolyzer or other system 1709 configured to produce chlorine gas and a metal hydroxide. In some implementations, the system 1701 is configured to produce hydrogen chloride or hydrochloric acid using chlorine, and the hydrogen chloride or hydrochloric acid may be delivered to the oxalic acid reactor 1707. In some implementations, the system 1701 is configured to utilize the hydroxide from the reactor 1709 within the formate reactor 1705.

[0361] As indicated, oxalic acid can be produced from metal formates by conversion to metal oxalates and subsequent acidification. In some implementations, the process may include the following steps: (1) producing a metal formate from carbon monoxide generated by a carbon dioxide electrolyzer, (2) producing a metal oxalate from the metal formate, e.g., by pyrolysis, and (3) producing oxalic acid by exposing the metal oxalate to an acid. The overall process can be implemented as a batch process. In some cases, at least the metal formate production operation and the metal oxalate production operation are performed in the same vessel. In some examples, the formate and / or oxalate production vessel may be a pressure vessel such as an autoclave. In some cases, the metal formate production vessel and / or the metal oxalate production vessel includes a mechanism for reducing the particle size of a solid reactant such as a metal hydroxide. In some examples, the production vessel includes a ball mill.

[0362] In some implementations, optionally during the reaction in an autoclave, the formate production reaction (e.g., the reaction of carbon monoxide and solid sodium hydroxide) has a residence time of at least about 15 to 60 minutes or about 20 to 40 minutes. In some cases, the solid hydroxide is contacted with carbon monoxide at a temperature of at least about 200°C (e.g., about 230 to 300°C) and / or at a pressure of at least about 2 bar (e.g., about 5 to 10 bar).

[0363] The operation of converting the metal formate to a metal oxalate may optionally be achieved by pyrolysis in the same reaction in which the metal formate was produced. In certain implementations, the chemical reaction is NaCOOH + CO → Na2C2O4. The reaction can be carried out in the presence of heat and sodium carbonate. In a batch process using a single reactor to produce the metal formate and the metal oxalate, the metal carbonate (e.g., sodium carbonate) can be the only input added to the reactor prior to the oxalate reaction.

[0364] In some embodiments, the metal oxalate formation reaction is carried out at a pressure of about 0.5 to 5 bar (e.g., substantially atmospheric pressure). In some embodiments, the metal oxalate formation reaction is carried out at a temperature of at least about 200 °C, or at least about 300 °C, or about 300 to 400 °C. In some embodiments, the residence time of the metal oxalate formation reaction is about 10 to 100 minutes, or about 20 to 40 minutes.

[0365] In some embodiments, at the start of the metal oxalate formation reaction, the pressure of the reactor is reduced to a low pressure (e.g., about 1 bar), and the reactor is heated until it reaches a temperature of at least about 300 °C (e.g., about 360 °C). During the thermal decomposition step, the metal hydroxide residue may continue to react with carbon monoxide, which increases the overall conversion of this reaction.

[0366] Solid metal carbonate can function as a catalyst. Furthermore, thermal decomposition of the metal oxalate salt into metal carbonate and carbon monoxide can be suppressed. In some embodiments, the pressure of the reactor is reduced to a low pressure (e.g., about 1 bar), and the reactor is heated until it reaches a temperature of at least about 300 °C (e.g., about 360 °C). During the thermal decomposition step, the metal hydroxide residue may continue to react with carbon monoxide, which increases the overall conversion of this reaction.

[0367] In certain embodiments, oxalic acid is formed from an acid such as metal oxalate and a hydrohalic acid. In certain embodiments, the acid is hydrochloric acid having a concentration of about 0.05 to 0.2 M. In some embodiments, the oxalic acid formation reaction is carried out in a crystallization apparatus such as a batch crystallization apparatus. In some embodiments, the reactor is configured to generate a recycle stream to enhance the efficiency of oxalic acid crystal separation.

[0368] In some embodiments, oxalic acid is produced at a temperature of about 20 to 100 °C, or about 50 to 100 °C. In some embodiments, the newly produced oxalic acid is cooled for about 10 to 60 minutes (e.g., to about 30 °C or below). In some embodiments, the pressure used during the oxalic acid formation reaction is about 0.5 to 2 bar (e.g., near atmospheric pressure). In certain embodiments, the reaction solution containing the metal oxalate is brought to a low pH, e.g., about 1 to 3 or simply about 1.

[0369] In some embodiments, the mass of water added to the metal oxalate is sufficient to complete the dissolution of the metal oxalate at the initial temperature. The dissolution process may occur while the temperature within the batch is, for example, about 80 °C. The crystallization process may occur when the temperature of the solution within the batch decreases, e.g., rapidly. This step may be enabled by using a bypass stream of the water initially used to dissolve the metal oxalate. That is, once the oxalic acid is completely dissolved, the water stream can avoid passing through the heater and flow directly to a heat exchanger placed in the batch. The process of forming and crystallizing oxalic acid can be achieved using a crystallization device such as the crystallization device 1731 shown in FIG. 17B.

[0370] The crystallized oxalic acid is optionally filtered and / or dried. In some embodiments, the oxalic acid is filtered using a pusher centrifuge. In some embodiments, the oxalic acid is dried using a fluidized bed dryer. In some embodiments, the filtered oxalic acid has a moisture content of about 20 to 30 mass%.

[0371] Figure 17B shows a process 1720 for forming oxalic acid from solid metal formate. As shown, process 1720 begins with the formation of metal formate in step 1721. This process includes the reaction of carbon monoxide from a carbon dioxide electrolyzer with a metal hydroxide. Any suitable process for producing metal formate may be used. Examples include all of the processes disclosed herein for producing metal formate, including the process described in connection with FIG. 9. After the metal formate is produced, in step 1723, the metal formate is converted to a metal oxalate. In the illustrated step, the conversion reaction from formate to oxalate is carried out in the presence of a metal carbonate such as sodium carbonate. In certain embodiments, one or both of steps 1721 and 1723 are carried out in a ball mill autoclave such as autoclave 1729 shown in the figure.

[0372] After the metal oxalate is formed in step 1723, the metal oxalate reacts with an acid such as a hydrohalic acid (e.g., hydrochloric acid) to form oxalic acid. See step 1725. In some embodiments, the reaction may be carried out in a crystallization device such as batch cooling crystallization device 1731 shown in the figure. The batch cooling crystallization device includes (1) a stirrer, (2) baffles, (3) a cooling jacket, (4) a jacket fluid inlet, (5) a jacket fluid outlet, and (6) an outlet valve. This arrangement facilitates both the formation and crystallization of oxalic acid.

[0373] After oxalic acid is formed and optionally crystallized in step 1723, the oxalic acid is optionally purified and / or dried. In one example, purification is achieved by filtration. In one example, drying is achieved in a fluidized bed.

[0374] FIG. 18 shows a system 1801 comprising a carbon dioxide electrolyzer 1803 configured to produce carbon monoxide and hydrogen. The system 1801 is configured to provide the carbon monoxide and hydrogen produced by the electrolyzer 1803 to a Fischer-Tropsch reactor or a gas fermentation reactor, as generally indicated by block 1805. Depending on the content of the product produced by the electrolyzer 1803 and the reaction taking place in the reactor 1805, additional hydrogen may be required to facilitate the reaction in the reactor 1805. For this purpose, the reactor 1805 may be configured to include an inlet for externally produced hydrogen gas. The reactor 1805 is configured to produce naphtha, and the system 1801 is configured to transfer the naphtha from the reactor 1805 to a naphtha cracking reactor 1807. The naphtha cracking reactor 1807 is configured to crack the naphtha in a manner that produces at least some propylene.

[0375] In certain embodiments, the carbon dioxide electrolyzer disposed upstream from the propylene production reactor is configured to operate in (a) the operating parameter regime of the hydrogen-rich product stream described herein, and / or (b) the operating parameter regime of the high reduction product to CO2 ratio described herein.

[0376] The system 1801 is configured to deliver propylene to an oxalic acid production reactor 1809 configured to receive nitric acid and optionally oxygen in addition to propylene. In some implementations, the system 1801 is configured to deliver excess oxygen from the electrolyzer 1803 to the reactor 1809. In certain embodiments, the reactor 1809 is configured to absorb propylene in nitric acid, add oxygen, remove nitrogen oxides, and heat the resulting mixture. The resulting process produces oxalic acid, and the system 1801 is configured to deliver it to a separation unit such as a crystallization device and separator unit 1811. In certain embodiments, the unit 1811 is configured to produce pure oxalic acid. The system 1801 may be configured to return nitric acid from the unit 1811 to the reactor 1809.

[0377] In one type of system 1801, propylene is produced from carbon monoxide and hydrogen generated in an electrolyzer by different routes. In this type, system 1801 is configured to deliver carbon monoxide and hydrogen to an alcohol synthesis reactor 1813 configured to produce methanol or other alcohols from a carbon dioxide electrolyzer 1803. In certain embodiments, reactor 1813 is configured to receive additional hydrogen from a source separate from electrolyzer 1803. In some implementations, reactor 1813 is configured to produce alcohol in a manner similar to the method of methanol synthesis reactor 1205 of FIG. 12.

[0378] Reactor 1801 is configured to transport the alcohol produced by alcohol synthesis reactor 1813 to a reactor 1815 from methanol to olefins configured to convert the alcohol to one or more olefins including propylene. System 1801 is configured to transport the propylene produced by reactor 1815 to oxalic acid synthesis reactor 1809. The reactor 1815 from methanol to olefins may be configured to convert an alcohol (e.g., methanol) to an olefin by a reaction including a network of chemical reactions in the presence of an acidic zeolite catalyst such as H-SAPO-34. The temperature and other parameters of the reaction may be adjusted to produce the desired product, which is propylene in system 1801. In certain embodiments, system 1801 is configured to operate reactor 1815 at a temperature of about 600-650 °C.

[0379] In various embodiments, a system that uses a carbon dioxide electrolyzer to produce carbon monoxide and hydrogen is configured to produce methanol from carbon monoxide and hydrogen by a method such as methanol synthesis reactor 1205 of FIG. 12. A related system may be configured to provide the obtained methanol to a formate synthesis reactor and be configured to perform the BASF and / or Kemira-Leonard process to produce methyl formate. Purification unit

[0380] To purify or otherwise concentrate the carbon monoxide produced by the carbon dioxide electrolyzer, various types of purification units may be used. By way of example, amine absorption units (e.g., used with a gas stream having a CO2 concentration of about 20% volume or less), CO2 adsorption units that utilize the acidity of CO2, CO adsorption units (e.g., use of copper compounds), CO / CO2 separation compositions such as molecular sieves and metal-organic frameworks, cryogenic systems (e.g., flash distillation systems), and membrane permeation units may be mentioned. In certain embodiments, the CO purification unit is configured to operate at a pressure of about 100 - 400 psia. The amine-based CO2 absorption unit may use an aqueous solution of an ethanolamine such as methyldiethanolamine MDEA, and optionally piperazine may be used to increase the absorption rate. The amine absorbent may be produced by applying heat. An exemplary unit uses an aqueous solution of about 30% (by weight) MDEA and 1% (by weight) piperazine.

[0381] Cryogenic systems function by cooling the gas mixture and then passing it through a fractionation column to separate the gases by boiling point. Multiple fractionation columns can be used in a single process to separate and deliver the purified components of the gas mixture.

[0382] The membrane purification process uses a membrane that retains the desired product gas but has high permeability to impurities in the gas stream. The membrane is packaged in a module into which the high-pressure gas mixture is input at the inlet. The membrane enables the desired product gas to be retained at high pressure and the unwanted impurity gas to be left in a separate low-pressure stream. For CO purification, the membrane retains CO but allows H2 and CO2 to pass through. The pure CO exits into the product stream. H2, CO2, and a small amount of CO remain in the waste stream. The low-pressure waste stream can be recompressed by a compressor and passed through another membrane stage to increase the recovery of the CO product. More than 99% pure CO can result from the membrane-based separation process.

[0383] A sorption process for removing CO may be used. The sorption process can use changes in pressure swing, temperature swing, vacuum swing, or other operating conditions (e.g., humidity swing). The sorbent is typically a solid or liquid that has a high affinity for the desired gas molecules under one extreme operating condition and a low affinity for the gas molecules under the other extreme operating condition. For example, a CO sorbent under high pressure conditions (e.g., about 300 kPa and about 40 - 60 °C) can capture CO (about 60 - 70 mol%) from a mixture of CO2 and H2. A system having four adsorption columns each containing 8 L of sorbent results in a flow rate of about 5 - 10 nm3 of 99 - 99.9% pure product CO. The use of larger columns or additional stages makes it possible to upgrade a CO-containing gas mixture of about 20 - 30% to a purity of about 99% or more. The sorbent for CO adsorption was developed by Kobe Steel using copper dopants in a solid matrix such as carbon, alumina, or silica.

[0384] Figure 19 shows a system 1901 for verifying a carbon monoxide stream containing carbon dioxide and, optionally, other components such as hydrogen. As shown, system 1901 includes an absorber 1903 configured to selectively remove carbon dioxide by contacting a carbon dioxide-containing gas with an absorption material such as ethanolamine. Absorber 1903 includes a gas inlet 1905 for receiving an inlet gas stream such as a gas stream from a carbon dioxide electrolyzer (not shown). Absorber 1903 also includes a gas outlet 1907 for discharging purified carbon monoxide.

[0385] Also, as shown, absorber 1903 includes an inlet 1911 for receiving purified sorbent and an outlet 1909 for discharging the loaded sorbent, e.g., sorbent containing a higher concentration of carbon dioxide than the sorbent entering absorber 1903.

[0386] System 1901 also includes a regenerator 1913 configured to remove carbon dioxide discharged by absorber 1903 and thereby generate a regenerated adsorbent material for reuse in absorber 1903. In the illustrated embodiment, the adsorbent is regenerated by heating and releases carbon dioxide. Heating is performed using preheater 1915 and reboiler 1917. Preheater 1915 is configured to receive the loaded adsorbent from outlet 1909 of absorber 1903 and deliver the preheated adsorbent to inlet 1919 of regenerator 1913. Preheater 1915 receives some heat from the lean adsorbent part way from regenerator 1913 to absorber 1903.

[0387] System 1901 is configured to transport lean adsorbent from outlet 1923 of regenerator 1913 to preheater 1915 using lean solvent pump 1921, where the adsorbent loses some of its heat. System 1901 is also configured to transport lean adsorbent from preheater 1915 to absorber inlet 1911 via trim cooler 1925. The trim cooler may be a water-cooled module.

[0388] As indicated, reboiler 1917 is configured to provide heat to regenerator 1913 to release carbon dioxide from the loaded adsorbent. As shown, reboiler 1917 receives a portion of the lean adsorbent from regenerator outlet 1923 and is included in a recirculation loop that returns the heated lean adsorbent to regenerator 1913 via inlet 1927.

[0389] Furthermore, system 1901 includes a recirculation compressor 1931 configured to compress carbon dioxide to a pressure suitable for entry, for example, into a carbon dioxide electrolyzer that produces an inlet carbon monoxide-containing stream.

[0390] System 1901 also includes a subsystem associated with the regenerator 1913 that uses a condenser 1933 to condense a portion of the adsorbent that may be contained in the carbon dioxide released from the regenerator 1913 through the gas outlet 1935. Note that the condenser 1933 is configured to condense the adsorbent and return it to the regenerator 1913 at the adsorbent inlet 1937.

[0391] In some cases, the CO purifier is a hybrid system having two different purification subsystems connected in series. For example, the CO purifier may have a cryogenic subsystem upstream of the adsorption subsystem. The hybrid system can be used, for example, in systems having a relatively low concentration of carbon monoxide such as an input CO stream of less than about 70 mol %.

[0392] FIG. 20 shows a hybrid carbon monoxide purification system 2001 having a cryogenic pretreatment subsystem 2003 and an adsorbent post-treatment subsystem 2005. The first stage of the system 2001 is a cryogenic subsystem 2003 configured to partially concentrate carbon monoxide. In certain embodiments, the cryogenic subsystem is configured to concentrate carbon monoxide to a level of at least about 70% by volume.

[0393] As shown, the cryogenic pretreatment subsystem 2003 is configured to supply the generated gas to a compressor 2007 configured to compress the gas to a defined pressure or density, for example, from a carbon dioxide electrolyzer (not shown). The system 2001 is configured to transport the compressed gas from the compressor 2007 to a chiller 2011 configured to reduce the temperature of the compressed gas. The chiller 2011 is coupled to a refrigeration system 2009 configured to remove sufficient heat from the chiller 2011 to maintain the compressed gas below a desired temperature.

[0394] Chiller 2011 is configured to cool the compressed gas to a reduced temperature. Chiller 2011 is also configured to discharge CO2 to a trim vaporizer 2013 that is configured to receive a CO2 slurry (to reduce the temperature) from a separator 2019 and discharge the vaporized carbon dioxide. The vaporized carbon dioxide from the trim vaporizer 2013 can be transported by the system 2001 to a recycle compressor 2015 that is configured to provide pressurized carbon dioxide suitable for supply to the cathode side of the carbon dioxide electrolyzer.

[0395] In the illustrated embodiment, the cryogenic subsystem 2003 is configured to receive the cooled and compressed gas from the output of the chiller 2011 to a Joule-Thomson valve 2017 that is configured to rapidly expand the compressed gas, thereby further cooling the gas. This action can sufficiently cool the gas to convert a portion of the gaseous CO2 to a liquid or solid or slurry. Nevertheless, the cryogenic gas is provided to a separator 2019 that is connected to the Joule-Thomson valve 2017 and has a carbon dioxide slurry outlet 2021 and a partially purified carbon monoxide gas outlet 2023. In certain embodiments, the partially purified carbon monoxide stream has a concentration of at least about 50 mol% or at least about 70 mol%.

[0396] In the illustrated embodiment, the system 2001 is configured to provide a partially purified carbon monoxide gas stream from the outlet 2023 to an adsorbent post-treatment subsystem 2005 that, in the illustrated embodiment, is configured similarly to the entire adsorbent system 1901. As shown, the subsystem 2005 has an absorber 2033 that is configured to absorb carbon dioxide from the partially purified CO stream and output a purified CO gas. The subsystem 2005 also has an adsorbent regenerator 2043. Other components of the subsystem 2005 include an adsorbent trim cooler, an adsorbent preheater, a lean adsorbent pump, an adsorbent reboiler, and an adsorbent condenser.

[0397] The descriptions of FIGS. 19 and 20, as well as other descriptions of the carbon monoxide purifier, describe cases including carbon dioxide removal, but the carbon monoxide purifier may be configured to additionally or alternatively remove other impurity gases such as sulfur-containing gases (e.g., sulfur oxides). Auxiliary hydrogen source

[0398] In certain embodiments, an integrated system that uses an oxide carbon electrolyzer includes an additional hydrogen source (beyond what is generated from the electrolyzer) or is configured to receive additional hydrogen from an external source. Examples of integrated systems that may use an additional hydrogen source include Fischer-Tropsch systems, polycarbonate production systems, ethylene glycol production systems, polyethylene terephthalate production systems, methanol, butanol, and / or other alcohol production systems, acetic acid production systems, isocyanate production systems, polyurethane production systems, and oxalic acid production systems. In certain embodiments, the additional hydrogen source is an electrolyzer such as a proton exchange membrane water electrolyzer. In some implementations, the electrolyzer shares the electrical infrastructure with the carbon dioxide reduction electrolyzer. In certain embodiments, the additional hydrogen source is (a) a unit or includes a unit configured to perform steam reforming, pyrolysis, and / or partial oxidation of methane, fuel oil, petroleum coke, and / or other fossil fuels, coal gasification, (b) steam methane reforming, (c) gasification, pyrolysis, and / or other high-temperature conversion of biomass, municipal solid waste, and / or other waste sources, (d) pressure swing adsorption of refinery waste streams, (e) separation of hydrogen by-products from industrial reactions such as molten salt chlorine production, and / or (f) dissociation of water by, for example, solar / thermal energy. In certain embodiments, the methane or other simple hydrocarbons used in one or more of these units are derived from biogas. Recovery of carbon dioxide from electrolyzer output

[0399] In many implementations, the product gas exiting the cathode of the carbon dioxide electrolyzer contains a significant proportion of unreacted carbon dioxide. For example, the product gas may contain about 10 - 70 mol% carbon dioxide. In certain embodiments, the system includes a carbon dioxide recovery unit configured to receive the product gas from the carbon dioxide electrolyzer and produce a concentrated carbon dioxide product, which may optionally be recycled to the electrolyzer. In one example, the carbon dioxide recovery unit includes a direct air carbon dioxide capture module as described elsewhere herein. It should be understood that the product gas from the carbon dioxide electrolyzer may contain a much higher concentration of carbon dioxide than air. Thus, the direct air capture unit used with the electrolyzer can have a modified configuration compared to the corresponding unit used for direct air capture. By way of example, the direct air capture unit can use temperature swing absorption, pressure swing absorption, or electric swing absorption.

[0400] FIG. 21A shows a system 2101 having an upstream DAC unit 2105 configured to concentrate carbon dioxide from air and a downstream DAC unit 2107 configured to remove unreacted CO2 from the product gas of the carbon dioxide electrolyzer 2103. The system 2101 is configured to combine the purified unreacted carbon dioxide with fresh carbon dioxide from the upstream DAC unit 2105 and introduce it into the electrolyzer 2103. The downstream DAC unit 2107 may be designed differently from the upstream DAC unit 2105 (e.g., amount of solvent, dimensions of the contactor) to account for the significantly different concentration of carbon dioxide in air compared to that in the product gas.

[0401] FIG. 21B shows a system 2111 having a DAC unit 2115 configured to capture carbon dioxide from air and separate unreacted carbon dioxide from the product gas of the electrolyzer 2113. The system 2111 is configured to supply the separated carbon dioxide to the electrolyzer 2113. The carbon dioxide dilute product and the carbon dioxide dilute air leave the DAC unit 2115 for downstream processing. Integration with the power grid

[0402] When the electrical energy source of the power grid cannot be directly controlled according to demand, various problems may occur. Solar, wind, and certain other non-combustion-based electrical energy sources are examples of sources where energy generation is decoupled from energy demand.

[0403] When renewable energy sources are connected to the power grid, fluctuations in wind speed or solar irradiance can reduce the amount of electricity available on the grid to the point where demand exceeds supply. This can cause the frequency of the power grid to drop, which can damage some electrical equipment and / or cause brownouts or blackouts.

[0404] To prevent this, responsive generators and consumers can be used in conjunction with renewable power sources. Such components facilitate a power grid load leveling system, are configured to draw responsive loads from the power grid, and can reduce power demand when necessary to prevent demand from approaching supply. This can provide frequency stabilization to the power grid and enable the power grid to operate with large amounts of renewable power.

[0405] Various techniques have been proposed for storing excess energy generated when supply exceeds demand. Examples include water storage tanks, batteries, and water electrolyzers. As an example, to use batteries to store excess energy on the power grid, a large number of high-capacity batteries are required to provide sufficient capacity to store the maximum excess energy that the energy source on the power grid can generate.

[0406] In comparison, a carbon dioxide electrolyzer can store excess energy in the form of a liquid or gas that is relatively easy to store. Also, compared to a water electrolyzer, a carbon dioxide reduction cell can be operated to produce a liquid product rather than a gaseous product. Liquid products may be easier to store, especially considering their relatively high density.

[0407] The products of the carbon dioxide electrolyzer can be used as fuel to generate electrical energy that is fed into the power grid during periods when demand may exceed supply. The electrolyzer products can be burned in a turbine or other mechanical power source and / or electrochemically consumed in a fuel cell to directly generate electricity. In certain embodiments, the carbon dioxide electrolyzer output, such as carbon monoxide or methanol, is stored for later use in a fuel cell and the electrical energy is injected directly back into the power grid. In certain embodiments, the fuel cell is a fuel cell configured to oxidize a carbon-containing reactant (e.g., natural gas), such as the solid oxide fuel cell from Bloom Energy of Sunnyvale, California.

[0408] The CO2 electrolysis products can be in the gas phase (e.g., CO, methane, ethylene) or the liquid phase (e.g., ethanol, methanol, ethylene glycol). The liquid products have the advantage of being easy to store for long periods. The gas phase products can be converted to liquid phase chemical compounds via a series of downstream processes such as gas fermentation or thermochemical reactions. The gas phase and liquid phase products can also be used to make solid materials. For example, CO is one of the inputs required to make polycarbonate or can react with potassium hydroxide to make potassium formate.

[0409] FIG. 22A shows a system 2201 that couples a carbon dioxide reduction electrolyzer 2203 to a power grid 2205 or other source of electrical energy. This system 2201 may be configured to operate in such a way as to store excess energy generated by the electrical system when the energy being generated exceeds the demand for electrical energy.

[0410] In system 2201, variable electrical energy source 2206 generates electrical energy that is optionally provided to power grid 2205. In the illustrated embodiment, energy source 2206 is coupled to power grid 2205 via an electrical wire having a rectifier 2207 and a transformer 2209. These and / or other electrical elements can be used to provide electrical energy from variable power source 2206 to power grid 2205 at an appropriate voltage and waveform.

[0411] In the illustrated embodiment, power grid 2205 is connected to a substantially constant electrical energy demand, at least compared to the variations of electrical energy supply 2206. In the illustrated embodiment, the demand is generally represented by element 2208, which represents one or more electrical energy consumers, such as residential and / or industrial consumers. In many applications, the energy consumers require electrical energy at a voltage and electrical waveform that can be generated by rectifier 2211 and transformer 2213. As shown, these elements are provided between power grid 2205 and demand 2208.

[0412] As shown, the illustrated embodiment provides an electrolytic carbon dioxide reduction cell or stack 2203 configured to consume excess electrical energy from a power grid or other electrical energy system by converting the excess energy into a chemical product of electrolytic carbon dioxide reduction. If this cell or stack 2203 receives electrical energy directly from the power grid, it may be necessary to rectify and convert the electrical energy by means of a rectifier 2217, a transformer 2219, and / or other electrical components.

[0413] The cathode side of electrolyzer 2203 receives carbon oxide reactants (carbon dioxide and / or carbon monoxide) via an inlet line. The carbon oxide may be provided by any one or more of a number of possible sources or feeds 2222 as described elsewhere herein.

[0414] The cathode is configured to produce products such as gaseous or liquid C1 compounds (e.g., carbon monoxide, methane, formaldehyde, or formic acid), or gaseous or liquid higher carbon compounds such as ethylene. Such products, along with other components, are removed from the electrolyzer 2203 via the outlet line. In certain embodiments, the outlet gas that may be humidified is provided to the gas separator 2225. This condenses water and / or one or more liquid products. Unreacted carbon monoxide and / or water may be returned to the carbon monoxide electrolyzer 2203 via line 2221. In certain embodiments, the carbon monoxide supplied to the electrolyzer is humidified to facilitate the reduction reaction.

[0415] In the illustrated embodiment, various optional components are provided downstream of the cathode side of the electrolyzer 2203. These include a demister 2227, a mass flow meter or controller 2229, a condensate trap 2231, a valve 2233, a gas tank 2235, a purification module 2237, and a product conversion system / reactor 2239. The resulting fuel, chemical, or other substance 2241 is appropriately stored or used. As described above, the chemical products of the electrolyzer 2203 store excess energy generated by a variable energy source. That energy may be recovered by converting the chemical energy stored in the products to thermal or electrical energy by combustion, fuel cell operation, etc. Alternatively, the energy is not recovered, at least not in the short term, and the electrochemically reduced products are converted to another useful commodity such as plastic.

[0416] Electrolyzer 2203 also includes an anode that receives anodic reactants such as water that is electrolytically oxidized at the anode. In the illustrated embodiment, the water supply source 2243 includes a water supply pump 2245, a demister 2247, a gas separation unit 2249 that also has an inlet for receiving products that may include water vapor and oxygen from the anode, a circulation pump 2251, a mass flow meter or controller, and one or more optional components such as an ion exchanger configured to remove ionic or other components of the anodic water that may be detrimental to the operation of the electrolyzer 2203, and is provided to the anode of the cell 2203 via these components. In a particular embodiment, the system 2201 includes a line that connects the anodic water recirculation loop to the carbon dioxide inlet line 2221 and humidifies the carbon dioxide delivered to the cathode side of the electrolyzer 2203.

[0417] The system 2201 may also have one or more components for removing oxygen or other products of the electrolyzer 2203. The anode side of the electrolyzer 2203 includes an outlet line 2207 configured to remove oxygen or other products from the electrolyzer. The outlet line 2207 is connected to the separation unit 2249. The oxygen and water in the anodic product stream may be separated from each other in the unit 2249, such that the water may be recirculated to the anode via the circulation pump 2251.

[0418] The system 2201 may be configured using components for removing oxygen or other gaseous products of the anode. In the illustrated embodiment, these components are a demister 2259 and a control valve 2261.

[0419] In these examples, and throughout the present disclosure, the described systems may be provided in a facility, plant, or building complex. In some embodiments, all or many of the reactors and units and / or modules of the system are provided in a common factory, plant, or complex. For example, a system for producing a particular material such as a polycarbonate polymer or a transportation fuel may comprise a carbon dioxide electrolyzer and one or more other reactors that utilize the electrolyzer's product and / or provide reactants to the electrolyzer, and the electrolyzer and the other reactors are provided in a single building or plant. In some cases, one or more system components are provided in the external environment. For example, even if the system is configured to provide carbon dioxide from a direct air capture unit directly to the electrolyzer, the direct air capture unit may be provided externally while the carbon dioxide electrolyzer is located inside a building.

[0420] Further examples of power grid management systems using electrolyzers are shown in FIGS. 22B and 22C. As shown, systems 2258 and 2259 use fuel cells configured to oxidize a carbon-containing reactant (e.g., methane or natural gas) such as a solid oxide fuel cell from Bloom Energy as described above. In the illustrated embodiment, carbon dioxide electrolyzer 2261 is coupled to a power grid or other electrical energy source 2263 such as a device configured to generate electricity from wind, sunlight, or other renewable energy sources. In addition to electrical energy, electrolyzer 2261 receives carbon dioxide and water as inputs. In certain embodiments, the carbon dioxide input is received, at least in part, from the output of fuel cell 2265. A compressor 2267 may be provided to compress such carbon dioxide prior to delivery to electrolyzer 2261.

[0421] As shown in the embodiment of FIG. 22B, the electrolyzer 2261 is configured to output carbon monoxide, and the system 2258 is configured to deliver the carbon monoxide to the fuel cell 2265. The system 2258 may be further configured to provide natural gas or other inputs to the fuel cell 2265, and the fuel cell is configured to generate electricity that can be provided to the power grid 2263. In some implementations, the system 2258 includes a steam methane reformer unit 2271 configured to produce hydrogen for input to the fuel cell 2265.

[0422] As shown in the embodiment of FIG. 22C, the electrolyzer 2261 is configured to output any one or more of various compounds including carbon monoxide, ethylene, methane, etc. In certain embodiments, one or more of these compounds are removed from the system 2259 for purposes potentially unrelated to electrical load leveling. For example, one or more of the compounds may be used as a feedstock for synthesizing compounds or polymers as described elsewhere in this specification.

[0423] In the illustrated embodiment, system 2259 is configured to deliver at least a portion of the output of electrolyzer 2261, optionally together with natural gas or other fuel from an external source, to fuel cell 2265. In certain embodiments, the output from electrolyzer 2261 is provided as synthetic natural gas or is converted to form such gas. System 2259 is configured to supply fuel cell 2265 with only synthetic natural gas or with other natural gas. In certain embodiments, system 2259 is configured to directly provide carbon monoxide and / or other outputs of electrolyzer 2261 to fuel cell 2265. In some implementations, system 2259 includes a steam methane reformer and / or a pressure swing adsorption unit 2273, or other purification units configured to process the gas before it enters fuel cell 2265. The steam methane reformer can convert a natural gas source and / or methane from the electrolyzer to hydrogen for delivery to fuel cell 2265. The pressure swing absorption unit can be used to remove carbon dioxide or other impurities from the inlet stream to fuel cell 2265. In some embodiments, another type of impurity removal unit is used.

[0424] In certain embodiments, such as those shown in FIGS. 22B and / or 22C, a pressure swing absorber or other gas purification unit is configured to separate a pure hydrogen stream from the gas mixture exiting the steam methane reformer. The tail gas from the reformer can contain CO, CO2, unreacted CH4, and some H2. The tail gas is supplied to the fuel cell, which then produces a stream of relatively pure CO2 (and water), which is then supplied to the CO2 electrolyzer. Some steam methane reformers are configured to utilize a water gas shift reaction after the reforming reaction to produce a product with a relatively high hydrogen concentration. The water gas shift reaction converts carbon monoxide and water (reactants) to carbon dioxide and hydrogen. In some implementations, the steam methane reformer is configured to produce a syngas mixture and may not perform a water gas shift step.

[0425] In some embodiments, as shown in FIGS. 22B and / or 22C, the steam methane reformer produces two CO2-containing streams, namely (1) the above-described tail gas coming from the reactor, and (2) the flue gas coming from the furnace used to heat the reactor tubes and generate steam for the reaction. The tail gas can contain a fairly high concentration of carbon dioxide (e.g., about 15% before separation and then about 50% after separation). The flue gas has a relatively low carbon dioxide concentration (e.g., a CO2 concentration of only about 3-5%). In various embodiments, about 2 / 3 of the carbon dioxide emissions are from the tail gas and about 1 / 3 are from the flue gas. In certain embodiments, the system is configured to capture emissions from (a) the tail gas before purification (e.g., using a pressure swing absorption unit), (b) the tail gas after purification, (c) the flue gas, or (d) the combined stream from the flue gas and the tail gas. In some embodiments, the system is configured to supply oxygen from the carbon dioxide electrolyzer to one or more furnaces of the steam methane reformer. In some embodiments, this achieves higher efficiency and / or results in a higher carbon dioxide concentration in the flue gas, thereby simplifying carbon dioxide capture.

[0426] The fuel cell 2265 is configured to output electricity that can be delivered to the power grid 2263. In certain embodiments, the system 2259 includes a carbon dioxide storage unit 2269 configured to store carbon dioxide output from the fuel cell 2265 before use by the electrolyzer 2261. Preparation of Synthesis Gas

[0427] The embodiments described in this section and / or illustrated in FIGS. 23A-23D relate to the production of mixtures of carbon monoxide and hydrogen. Some of these mixtures may be called synthesis gas. The embodiments described herein relate to methods and systems configured to receive a mixture of carbon monoxide, carbon dioxide, and hydrogen and reform the mixture to produce a mixture of carbon monoxide and hydrogen having a specific composition. In some cases, the input is a gaseous mixture obtained from a carbon dioxide electrolyzer, such as one of the carbon dioxide electrolyzers described herein.

[0428] As described herein, the mixture of carbon monoxide and hydrogen produced can have various uses. This can be used to produce naphtha, or other liquid hydrocarbon compositions that can be produced by the Fischer-Tropsch process (see, for example, the descriptions of FIGS. 6A and 6B). It can also be used as an input to a gas fermentation reactor (see, for example, FIGS. 4 and 5). It can also be used to produce any of various chemicals such as alcohol (see, for example, FIGS. 15 and 18) and / or polyol (see, for example, FIGS. 14 and 16).

[0429] Various embodiments for producing a mixture of carbon monoxide and hydrogen can use a carbon monoxide separator system as described in connection with FIGS. 19 and 20.

[0430] The mixture of carbon monoxide and hydrogen can be produced by directly removing either CO or CO2 from an input stream. In embodiments where the mixture is produced by directly separating CO from the input stream, hydrogen may be added to the purified CO stream downstream of the CO purification step. For example, purified hydrogen may be prepared by separating it from CO2 in a step downstream of the CO separation step.

[0431] Figure 23A shows a general scheme for generating a mixture of carbon monoxide and hydrogen in a process for directly separating carbon monoxide from an input stream. The input stream that can be provided from the cathode outlet of the carbon dioxide electrolyzer contains carbon dioxide, carbon monoxide, hydrogen, and optionally other components such as small amounts of water and / or hydrocarbons. The input stream is fed to one or more separation elements 2303 configured to produce one stream 2305 containing purified carbon monoxide and another stream 2307 containing a mixture of carbon dioxide and hydrogen. The element 2303 can include, for example, a CO absorption element and / or a CO absorption and stripping subsystem such as a pressure swing subsystem or a temperature swing subsystem. The stream 2307 is fed to one or more elements 2309 configured to separate hydrogen from carbon dioxide. In some embodiments, the element 2309 includes, for example, a membrane separator that blocks the passage of carbon dioxide while allowing the passage of hydrogen. During operation, the element 2309 produces a hydrogen stream 2311 and a carbon dioxide stream 2313. The hydrogen stream 2311 can be combined with the carbon monoxide stream 2305 to produce a stream 2315 containing a mixture of carbon monoxide and hydrogen. The carbon dioxide stream 2313 can optionally be recycled to the carbon dioxide electrolyzer.

[0432] In some embodiments, the process for generating a mixture of hydrogen and carbon monoxide can be characterized by the following steps. 1. Separation of CO from the mixture, for example, by ionic liquid absorption in a pressure swing absorption process 2. Separating H2 from CO2, for example, through a membrane 3. Mixing H2 and CO

[0433] In certain embodiments for producing purified carbon monoxide directly from an input stream, an ionic liquid is used to remove carbon monoxide from the input gas stream. During separation, the ionic liquid contacts the input gas and selectively absorbs carbon monoxide while allowing most of the hydrogen and carbon dioxide to pass through (non-dissolved or non-absorbed). In some embodiments, the input stream contacts the ionic liquid within an absorption column. After contact, the input stream, i.e., the carbon monoxide-rich stream of the ionic liquid, is fed to a stripper operating under conditions to remove carbon monoxide from the ionic liquid. The resulting lean stream of the ionic liquid may be recycled back to a component that selectively absorbs carbon monoxide.

[0434] Ionic liquids suitable for separating carbon monoxide preferentially absorb carbon monoxide without substantially absorbing carbon dioxide and / or hydrogen. Other characteristics may include low cost, low vapor pressure (e.g., no generation of volatile organic compounds during use), low kinematic viscosity (e.g., less than about 150 cSt), moderate absorption conditions (e.g., a temperature of about 0°C to 20°C, a pressure of about 25 bar or less (e.g., about 17 bar)), moderate stripping conditions (e.g., a temperature of about 0 to 100°C or less, a pressure of about 5 bar or less), and / or low toxicity. Examples of such ionic liquids include 1-hexyl-3-methylimidazolium chloride (including cuprous chloride).

[0435] In some embodiments, a mixture of carbon monoxide and hydrogen is produced by directly removing carbon dioxide from an input stream containing carbon monoxide, hydrogen, and carbon dioxide. In such embodiments, one output stream of the separation contains the desired mixture of hydrogen and carbon monoxide.

[0436] Figure 23B shows a general scheme for producing a mixture of carbon monoxide and hydrogen in a process for directly separating carbon dioxide from an input stream. The input stream that can be provided from the cathode outlet of the carbon dioxide electrolyzer contains carbon dioxide, carbon monoxide, hydrogen, and optionally other components such as small amounts of water and / or hydrocarbons. The input stream is fed to one or more elements 2302 configured to produce a carbon dioxide stream 2304 containing purified carbon dioxide and another stream 2306 containing a mixture of carbon monoxide and hydrogen. The element 2302 can include, for example, a carbon dioxide absorption element or a subsystem including a carbon dioxide absorption element and a stripping element, such as a pressure swing or temperature swing subsystem. The carbon dioxide stream may optionally be recycled to the carbon dioxide electrolyzer.

[0437] In a particular embodiment for directly removing carbon dioxide from the input stream, an ionic liquid is used to remove carbon dioxide from the input gas stream. The ionic liquid contacts the input gas and selectively absorbs carbon dioxide while allowing most of the hydrogen and carbon monoxide to pass through. In some embodiments, the input stream contacts the ionic liquid within an absorption column. After contact, the input stream, i.e., the carbon dioxide-rich stream of the ionic liquid, is fed to a stripper operating under conditions to remove carbon dioxide from the ionic liquid. The resulting lean stream of the ionic liquid may be recycled back to the component that selectively absorbs carbon dioxide for reuse.

[0438] Ionic liquids suitable for separating carbon dioxide preferentially absorb carbon dioxide without substantially absorbing carbon monoxide and / or hydrogen. Other characteristics can include low cost, low vapor pressure (e.g., no generation of volatile organic compounds during use), low kinematic viscosity (e.g., less than about 150 cSt), moderate absorption conditions (e.g., a temperature of about 15 °C or higher, a pressure of about 50 bar or lower), moderate stripping conditions (e.g., a temperature of about 0 to 100 °C or lower, a pressure of about 5 bar or lower), and / or low toxicity. Examples of such ionic liquids include 1-butyl-3-methylimidazolium hexafluorophosphate [bmim][PF6].

[0439] In some implementations, a system configured to produce a mixture of carbon monoxide and hydrogen does not include a component configured to cool the input or effluent to less than about 30 °C or less than about 20 °C or less than about 10 °C. For example, the system does not include a compressor that cools the inlet gas to less than about 20 °C. CO Separation Example

[0440] As described, the output from the carbon dioxide electrolyzer contains carbon monoxide and hydrogen. This electrolyzer output gas functions as an input to a separation system that produces a mixture of carbon monoxide and hydrogen. In some embodiments, carbon monoxide is recovered from the gas stream by a pressure swing absorption process using an ionic liquid such as 1-hexyl-3-methylimidazolium chloride (CuCl).

[0441] The pressurized inlet gas stream enters the bottom of the absorption column, and the absorbent is sprayed from the top of the column. The packing layer can enhance the contact between the gas phase and the liquid phase and promote carbon monoxide absorption. The remaining gas compounds (CO2, H2, a small amount of residual CO) exit from the top of the column, while the CO-enriched liquid phase exits from the bottom of the column.

[0442] In certain embodiments, the absorbent comprises a mixture of CuCl and an ionic liquid (e.g., about 50 mol% CuCl and 50 mol% ionic liquid 1-hexyl-3-methylimidazolium chloride). The inlet molar amount of CuCl may be about 1.5 to 2.5 times (e.g., about 1.9 times) greater than the amount of carbon monoxide to be absorbed.

[0443] After being heated in a heat exchanger, the CO-rich liquid phase enters a stripping column. The pressure of the stripping column can be less than 2 bar (e.g., approximately atmospheric), and the temperature can be relatively low at the top (e.g., about 30 °C) and relatively high at the bottom (e.g., about 60 °C). Under these conditions, most of the CO may evaporate. The liquid phase flows to the bottom of the column, where a portion (e.g., about 15%) is recycled back to the column using, for example, a total reboiler, and the remainder (e.g., about 85%) flows back to the absorption column. In certain embodiments, the incoming boiling liquid from the reboiler transfers heat to the incoming stream.

[0444] Refer to the following CO absorption stripping components of the system 2321 shown in FIG. 23C. The inlet stream 2322 is compressed and cooled by a compressor 2323 operating in conjunction with a chiller 2324. The compressed and cooled inlet gas enters the bottom of an absorption column 2325 where CO in the inlet stream is selectively absorbed by a liquid. The CO-rich liquid phase is heated in a heat exchanger 2326 (stream 5) and then enters a stripping column 2327 (stream 6). The pressure of the stripping column may be atmospheric pressure, and the temperature may range, for example, from 30 °C at the top to 60 °C at the bottom. The solubility of CO in the absorbent at these operating parameters is low, causing most of the CO to evaporate. The liquid phase flows to the bottom of column 2327 (stream 8), where, for example, about 15% is recycled back to column 2327 using a reboiler 2328 (streams 9 / 10), while the remainder, for example 85%, flows back to absorption column 2325 (streams 11 - 16).

[0445] The gas streams of CO2 and H2 exit the top of the absorption column 2325. CO2 and H2 can be separated by various techniques. In some embodiments, they are separated using a membrane filter 2330. In some embodiments, the membrane filter is a Polaris (trademark) filter manufactured by Membrane Technology And Research Corporation Inc., Newark, California. Such a membrane can have a high permeability to CO2 compared to H2 obtained as the retained fluid. In some examples, the membrane filter operates at a pressure of about 5 to 15 bar (e.g., a pressure of about 9 bar) and / or a temperature of about 0 to 20 °C (e.g., about 5 °C). CO2 Separation Example

[0446] Similar to the embodiments of CO absorption, carbon monoxide and hydrogen are included in the output from the carbon dioxide electrolyzer. This electrolyzer output gas functions as an input to a separation system that produces a mixture of carbon monoxide and hydrogen. In certain implementations, CO2 is directly absorbed from the input stream using an ionic liquid such as [bmim][PF6]. The absorption can be carried out at any of various pressures (e.g., about 10 to 60 bar). In some implementations, the pressure of the inlet gas stream can be about 8 to 16 bar. For this purpose, the system may utilize additional compression by a compressor.

[0447] In some embodiments, the CO2 absorption column is operated at a relatively high pressure such as about 20 to 60 bar or about 40 to 50 bar (e.g., about 44 bar). At this pressure and 25 °C, the solubility of CO2 is about 0.2 moles in 1 mole of ionic liquid. In some embodiments, the CO2 absorption column is operated at a relatively low pressure such as about 1 to 20 bar. In certain embodiments, the absorption column is operated at a temperature of about 20 to 80 °C, or about 20 to 30 °C, or about 40 to 60 °C.

[0448] Figure 23D shows an exemplary CO2 absorption-stripping system 2351 for processing the output of a carbon dioxide electrolyzer to produce a mixture of carbon monoxide and hydrogen. The inlet stream is compressed by compressor 2353. The compressed inlet gas (stream (2)) enters the bottom of the absorption column 2355 where CO2 in the inlet stream is selectively absorbed by a liquid (e.g., an ionic liquid). The flow rate of the absorbent may be adjusted or controlled by monitoring the molar flow rate of carbon dioxide in the inlet stream. The absorbent within the column may be maintained at a particular level. Since the absorbent recirculates through the system, it may not accumulate at the bottom of the column. A controller may be used to sense the amount of absorbent at the bottom of the column and adjust or maintain the absorbent level at a desired level within the column.

[0449] In certain embodiments, the CO2 absorption column operates under the following conditions, namely, Partial pressure of CO2: 13 bar Temperature: 28.1 °C Liquid phase viscosity: 110 cP Gas phase viscosity: 0.017 cP Diffusivity of CO2 in the absorbent: 500 mm 2 / s Column diameter: 0.5 m Column height: 12.4 m Pressure drop: 0.012 bar Cf (packing constant): 170 Total specific surface area of packing: 108 m 2 / m 3 The nominal packing diameter is 0.005 m.

[0450] The gas streams of CO and H2 (stream (3)) exit the top of the absorption column 2355. This stream may be used as syngas.

[0451] The CO2-enriched liquid phase from the absorption column 2355 enters the stripping column 2357 (stream 6) after being heated in the heat exchanger 2356 (stream 5). The solubility of CO2 in the absorbent in the stripper's operating parameters is low, evaporating most of the CO2. The liquid phase flows to the bottom of the stripping column 2357 (stream 8), where a portion of it is recycled back to the column 2357 using the reboiler 2358 (streams 9 / 10), and the remaining liquid flows back to the absorption column 2355 (streams 11 - 16). In certain embodiments, about 4 - 5% of the bottom liquid at the bottom of the absorption column is recycled back to the column using, for example, a reboiler.

[0452] In certain embodiments, the CO2 stripping column is operated at a pressure of about 0.5 - 5 bar (e.g., about 1 bar), the inlet CO2-rich absorbent stream has a temperature of about 40 - 60 °C, and the bottom of the column has a temperature of about 60 - 80 °C. Ethylene purification

[0453] The embodiments described in this section relate to the production of ethylene. The embodiments described herein relate to methods and systems configured to receive a mixture containing ethylene and reform the mixture to produce purified ethylene. In some cases, the input is a gaseous mixture obtained from a carbon dioxide electrolyzer, such as one of the carbon dioxide electrolyzers described herein.

[0454] In some implementations, the ethylene purification system is configured to produce relatively pure ethylene without necessarily producing a relatively pure stream of any other components produced by the electrolyzer. In some implementations, the ethylene purification system is configured to produce relatively pure ethylene along with a relatively pure stream of one or more other components, such as hydrogen, carbon monoxide, carbon dioxide, methane, ethanol, or any combination thereof.

[0455] In various embodiments, an ethylene purification system includes (a) absorbing and separating carbon dioxide, (b) separating ethylene from one or more other components by membrane filtration, (c) fractional distillation to separate ethylene and methane, (d) chemically converting methane to ethylene, and (e) one or more components or subsystems for any combination of (a)-(d). In some embodiments, an amine or ionic liquid is used to absorb carbon dioxide. In some embodiments, the membrane filtration component is configured to separate carbon monoxide and hydrogen (optionally with methane) from ethylene.

[0456] Ethylene produced as described herein can have various uses. For example, it can be used to produce ethylene oxide (see, e.g., the descriptions of FIGS. 10A, 11, and 14), and optionally reaction products of ethylene oxide such as monoethylene glycol and polyethylene glycol.

[0457] In certain embodiments described in this section, the input gas includes ethylene, typically some methane and unreacted carbon dioxide. Other components that may be present include hydrogen, carbon monoxide, water, ethanol, and any combination thereof.

[0458] As an example, the inlet stream to the ethylene purification system may have the following molar % composition: hydrogen (4.75%), methane (23.72%), carbon monoxide (0%), carbon dioxide (50.73%), ethylene (9.49%), ethyl alcohol (4.75%), and water (6.57%). Such a composition can be produced by a carbon dioxide electrolyzer. Route 1: Cryogenic distillation for ethylene separation

[0459] In some embodiments, ethylene is separated from other components by a route that includes absorption of carbon dioxide and subsequent fractional distillation to remove hydrogen, carbon monoxide, and / or methane to produce purified ethylene. As an example, the process can include the following series of steps. Step 1: Condensation of the liquid product to remove ethanol and water Step 2: CO2 removal Step 3: Distillation - removal of hydrogen, CO, and methane Step 3 (alternate or optional): Methane conversion to ethylene by oxidative coupling of methane

[0460] The separation process for removing water and ethanol from the input stream can be carried out in various ways. In some embodiments, to further remove water and ethanol, it is carried out in a two-step process of condensation followed by a molecular sieve absorption process. In some embodiments, the condensation of water and ethanol is achieved using compression of the input stream. In some embodiments, the condensation of water and ethyl alcohol is achieved using an absorption column, such as a countercurrent column, during production with water. The column optionally contains a catalyst. In some implementation forms, thermal equilibrium is reached within the column and both outlet streams have the same temperature (e.g., about 25 - 50 °C).

[0461] In some cases, amines such as diethanolamine, monoethanolamine, dimethylamine, piperazine, 2-aminopropanol, diisopropanolamine, aminoethoxyethanol, and / or methyldiethanolamine are used, or carbon dioxide is removed from the gas stream by an ionic liquid. In some implementation forms, the concentration of the selected amine is at least about 10 times higher than the concentration of carbon dioxide. In some embodiments, the amine-containing liquid has about 50 - 80 mole percent amine in an aqueous solution.

[0462] Exemplary operating conditions for an amine-based carbon dioxide removal process are Absorber: about 35 to 50 °C, absolute pressure about 5 to 205 atm; Regenerator: about 100 to 126 °C and an absolute pressure of about 1.4 to 1.7 atm at the bottom of the column.

[0463] In some embodiments, a temperature difference of about 5 °C or more is maintained between the lean amine and the sour gas. If the temperature difference is closer, hydrocarbon condensation may occur.

[0464] To demethanize an ethylene- and methane-containing stream, various techniques can be used. In some embodiments, a cryogenic distillation process is performed. See, for example, U.S. Patent No. 3,902,329 (King III et al.), which is incorporated herein by reference in its entirety. In some embodiments, the cryogenic distillation is performed at a temperature of about -90 °C or less.

[0465] In some embodiments, an ethylene / methane gas mixture is pressurized in a compressor (e.g., to a pressure of about 100 bar and an outlet temperature of about 15 °C). The gas mixture is cooled with chilled water. Then, for example, by throttling the flow of the compressed gas mixture with a throttle valve, the outlet gas may be substantially cooled (e.g., to a temperature of about -100 °C).

[0466] An exemplary process for the separation of methane from ethylene by cryogenic distillation is presented below, namely, 20 to 30 plates or more, Temperature of the plate where condensate is returned to the column: -90 °C to -105 °C, Pressure: 25 to 40 bar or more 98% efficiency.

[0467] In some examples, a cryogenic distillation column has the following design parameters, namely, Diameter: 0.085 m Height 6.8 m Trays: 17 Tray efficiency: 80% Reflux ratio: 1.129 Vapor linear velocity: 3 m / s Separation efficiency: 98%.

[0468] In some embodiments, the process removes hydrogen from ethylene (and optionally other components) via membrane separation.

[0469] In some embodiments, the process uses oxidative coupling of methane (OCM) to ethylene. OCM may be performed on methane-containing vapor after separation of methane and ethylene. This process may produce ethane, CO, H2, and CO2 as undesirable by-products. In addition to the temperature of the reaction, an important parameter is the amount of oxygen reacting with methane.

[0470] OCM may include some or all of the reactions presented below. For example, see Bhatia, Subhash & Thien, Chua & Mohamed, Abdul (2009), Oxidative Coupling of Methane (OCM) in Catalytic Membrane Reactors and Its Performance Comparison with Other Catalytic Reactors, Chemical Engineering Journal - CHEM ENG J. 148. 525 - 532. 10.1016 / j.cej.2009.01.008, which is hereby incorporated by reference in its entirety. [Table 2]

[0471] The yield of this process depends on the reaction conditions and the oxygen ratio. The amounts of methane and oxygen can be selected, for example, to facilitate the reactions in steps 2 and 5. In certain embodiments, this process is performed in a catalytic membrane reactor comprising a disk-shaped planar BSCF membrane.

[0472] Examples of the design and operation of the OCM reactor are shown in the following table of Handbook of Membrane Reactors: Reactor Types And Industrial Applications, 2013 by X. Tan, K. Li, which is hereby incorporated by reference in its entirety.

Table 3

[0473] In various embodiments, the OCM temperature is within the range where steam cracking of ethane produces ethylene. Analysis indicates that operating in the temperature range of 850 °C to 950 °C and a steam-to-hydrocarbon ratio of 0.3 to 0.5 results in good ethylene yields while minimizing by-products. In some embodiments, the OCM cracking reaction is carried out at high pressure (e.g., about 2 to 2.5 bar) in a tubular reactor.

[0474] In certain embodiments, about 0.3 of methane is converted to ethylene (moles), which can be approximately the amount of ethylene present in a typical inlet feed. Thus, by utilizing the above process, the amount of ethylene produced is nearly doubled.

[0475] In some embodiments, the process includes a step of separating steam and hydrogen from the ethylene formed. An absorption countercurrent column may be used for this step. As an example, the process conditions can include a pressure of about 5 to 50 bar (e.g., about 10 bar) and a temperature of about 150 to 500 °C (e.g., about 300 °C). In some embodiments, an absorption column with an appropriate volume of water flow separates nearly 100% of the water and hydrogen from the gas. Route 2: Use of a membrane for ethylene separation

[0476] In some embodiments, ethylene is separated from other components by a route that includes membrane separation of the gas stream to produce an ethylene-rich stream. As an example, such a process can include the following series of steps, namely Process 1: Condensation of the liquid product to remove ethanol and water Process 2: CO2 removal by amine treatment Process 3: CO + H2 removal by membrane separation Process 4: Ethylene separation from methane using membrane separation

[0477] In some implementations, Process 1 and Process 2 are carried out in the same manner as the above-described route using cryogenic distillation. Process 3 and Process 4 are carried out using membranes designed or configured to separate gaseous components from each other. In certain embodiments, suitable membranes are provided by Membrane Technology & Research, Inc., Newark, California.

[0478] The design may include a compression stage to remove ethanol and water prior to membrane separation. Membrane stage for removing non-hydrocarbons

[0479] The first membrane can separate nearly 100% of H2, CO2, ethanol, and water. It may not significantly separate CH4 from C2H4.

[0480] As an example, starting from an initial gas mixture of 50.7 mol% CO2, 23.7 mol% CH4, and 9.5 mol% C2H4, the resulting product stream contains 64.7 mol% CH4 and 30.5 mol% C2H4, and most of the remaining gas is CO2 (3.7%). To reduce the loss of the CH4 / C2H4 mixture in the permeate stream, a two-stage separation design can be employed.

[0481] The pressure is 10 bar and the inlet stream is at a temperature of 30°C. The cooling process takes place within the membrane and the temperature of the residue is -2.1°C. A 10% loss is satisfactory.

[0482] In some embodiments, the membrane is a hollow fiber membrane comprising polypropylene (PP), polyethylene (PE), polytetrafluoroethylene (PTFE), PVDF, polysulfone (PS), polyetherimide (PEI), or any combination thereof. In some embodiments, the membrane has a porosity of about 50 - 70% (e.g., about 60%). In some embodiments, the membrane has an average pore size of about 2 - 3 μm. Ethylene / methane separation

[0483] In certain embodiments, the methane-ethylene separation membrane comprises a metal-organic membrane for separation at room temperature. In certain embodiments, the methane-ethylene membrane separation has an adsorption selectivity of 12 to 20 at 296K. An adsorption selectivity of 20 represents a separation process efficiency of 95%.

[0484] The membrane may comprise a microporous metal-organic framework Zn4L(DMA)4 (UTSA-33, H8L = 1,2,4,5-tetra(5-isophthalic acid)benzene, DMA = N,N'-dimethylacetamide) having small pores of about 4.8 to 6.5 Å (He, Yabing et al. "A microporous metal-organic framework for highly selective separation of acetylene, ethylene, and ethane from methane at room temperature", Chemistry - A European Journal 18.2 (2012): 613, which is hereby incorporated by reference in its entirety).

[0485] In some implementations, after methane separation, the methane is subjected to OCM to increase the yield of ethylene.

[0486] Route 3: Use of a filtration membrane and cryogenic distillation for ethylene separation

[0487] In certain embodiments, a membrane filter is used to separate methane and ethylene from other components such as hydrogen, carbon monoxide, and carbon dioxide. In some embodiments, the methane-ethylene mixture is then separated into relatively pure ethylene and methane streams by cryogenic distillation. In some implementations, a separate membrane filter is used to separate hydrogen from carbon monoxide, carbon dioxide, and optionally other components.

[0488] In some implementations, the process may include the following steps, namely Step 1: Compression of the gas stream to enable condensation of ethyl alcohol and water Step 2: Removal of water and ethyl alcohol in a countercurrent absorption column using a catalyst Step 3: Membrane filtration to separate methane and ethylene from other gases such as carbon dioxide, carbon monoxide, and hydrogen Step 4 (optional): Membrane filtration to separate hydrogen from carbon monoxide and carbon dioxide Step 5: Cryogenic distillation to separate methane and ethane (optionally using a methane output stream cooled as a cooling utility) Oxygen generation

[0489] The carbon dioxide electrolyzer anode can produce oxygen from water. Oxygen can be used in any of various integration schemes of the electrolyzer. In some cases, oxygen can be used in a combustion reaction with fuel. In some cases, oxygen can be compressed and stored for later use. In certain embodiments, the compressed oxygen is cooled and then passed through a throttle valve to liquefy the oxygen. The cooling may be achieved using a Freon-type cooler. In some cases, the oxygen stream is first cooled using a brine cooler (e.g., using CaCl2 brine). For example, at 40 bar and -120 °C, oxygen becomes liquid. In some implementations, the oxygen stream is cooled to about -70 °C or lower. Embodiments of the controller

[0490] In embodiments that use a controller or other logic to control the operation of one or more reactors, pumps, separators, and / or other components of a system, the controller or logic can use program instructions, such as executable instructions on a computer-readable medium. The instructions may be executed by computer-executable components, such as those integrated with a communication system. The computer-readable medium may be stored on any suitable computer-readable medium, such as RAM, ROM, flash memory, EEPROM, optical devices (CD or DVD), hard drives, floppy drives, or any other suitable device. The computer-executable components may optionally be a processor, but the instructions may alternatively or additionally be executed by any suitable dedicated hardware device.

[0491] Although omitted for brevity, embodiments of the system and / or method can include all combinations and permutations of various system components and various method processes, and one or more examples of the methods and / or processes described herein can be executed asynchronously (e.g., sequentially), simultaneously (e.g., in parallel), or in any other suitable order by one or more examples of the systems, elements, and / or entities described herein and / or by using those examples.

[0492] The figure shows the architecture, functions, and operations of possible implementations of a system, method, and computer program product according to the disclosed embodiments, exemplary configurations, and their variations. In this regard, each block in the flowchart or block diagram can represent a module, segment, step, or part of code that includes one or more executable instructions for implementing the specified logical function. It should also be noted that in some alternative implementations, the functions described in the blocks may be performed in an order different from that shown in the figure. For example, two blocks shown consecutively may actually be executed substantially simultaneously, or the blocks may sometimes be executed in the reverse order depending on the related functions. It should also be noted that each block of the block diagram and / or flowchart diagram, and combinations of blocks of the block diagram and / or flowchart diagram, can be implemented by a dedicated hardware-based system for performing the specified function or operation, or by a combination of dedicated hardware and computer instructions.

[0493] Those skilled in the art will recognize that, without departing from the scope of the present disclosure as defined in the following claims, modifications and changes can be made to the disclosed embodiments of the present disclosure, as will be recognized from the foregoing detailed description, drawings, and claims. [Considered Items] [Item 1] A system for producing a polycarbonate polymer, (a) A carbon dioxide reduction electrolyzer having a membrane electrode assembly including one or more ion-conductive polymer layers and a cathode catalyst for promoting the chemical reduction of carbon dioxide to carbon monoxide; (b) A plurality of intermediate reactors collectively configured to receive the carbon monoxide produced by the carbon dioxide reduction electrolyzer and produce one or more intermediate chemicals; and (c) A polycarbonate synthesis reactor configured to receive the one or more intermediate chemicals or one or more derivatives of the one or more intermediate chemicals and synthesize a polycarbonate polymer A system comprising [Item 2] wherein the plurality of intermediate reactors include a bisphenol A synthesis reactor configured to produce bisphenol A, wherein the one or more intermediate chemicals include the bisphenol A, The system according to item 1, wherein the system is configured to transport the bisphenol A from the bisphenol A synthesis reactor to the polycarbonate synthesis reactor. [Item 3] The system according to item 1, wherein the plurality of intermediate reactors include a gas fermentation reactor. [Item 4] The system according to item 3, wherein the gas fermentation reactor is configured to produce a product selected from the group consisting of naphtha, phenol, and acetone. [Item 5] The system according to item 1, wherein the plurality of intermediate reactors include a Fischer-Tropsch reactor configured to produce naphtha. [Item 6] The system according to item 1, wherein the plurality of intermediate reactors include a naphtha cracking reactor configured to produce unsaturated hydrocarbons from naphtha. [Item 7] The plurality of intermediate reactors further include a Fischer-Tropsch reactor and / or a gas fermentation reactor configured to produce the naphtha, The system according to item 6, wherein the system is configured to transport the naphtha from the Fischer-Tropsch reactor and / or the gas fermentation reactor to the naphtha cracking reactor. [Item 8] The plurality of intermediate reactors further include a cumene process reactor configured to produce phenol and acetone from the unsaturated hydrocarbons, The system according to item 6, wherein the system is configured to transport the unsaturated hydrocarbons from the Fischer-Tropsch reactor and / or the gas fermentation reactor to the cumene process reactor. [Item 9] The plurality of intermediate reactors include a bisphenol A synthesis reactor configured to produce bisphenol A, wherein the one or more intermediate chemical substances include the bisphenol A, The system according to item 8, wherein the system is configured to transport the phenol and the acetone produced by the cumene process reactor to the bisphenol A synthesis reactor. [Item 10] The system according to item 1, wherein the plurality of intermediate reactors include a phosgene production reactor configured to produce phosgene from carbon monoxide and chlorine. [Item 11] The system according to item 10, further comprising a chlor-alkali cell configured to produce the chlorine. [Item 12] The system according to item 11, wherein the chlor-alkali cell and the carbon dioxide reduction electrolytic cell are configured to share electricity from a common power source. [Item 13] The system according to item 11, wherein oxygen produced by the carbon dioxide reduction electrolytic cell is delivered to the chlor-alkali cell. [Item 14] The system according to item 1, further comprising a carbon dioxide recirculation system configured to deliver carbon dioxide from at least one of the plurality of intermediate reactors to the carbon dioxide reduction electrolytic cell. [Item 15] The system according to item 14, wherein at least one of the plurality of intermediate reactors has a Fischer-Tropsch reactor or a gas fermentation reactor. [Item 16] The system according to item 1, wherein the carbon dioxide reduction electrolytic cell is configured to produce hydrogen and the carbon monoxide at a molar ratio of at least about 0.5:1. [Item 17] The system according to item 1, wherein the carbon dioxide reduction electrolytic cell, the plurality of intermediate reactors, and the polycarbonate synthesis reactor are arranged in a single plant. [Item 18] A method for producing a polycarbonate polymer, comprising: (a) reducing carbon dioxide to carbon monoxide in a carbon dioxide reduction electrolytic cell having a membrane electrode assembly including one or more ion conductive polymer layers and a cathode catalyst for promoting the chemical reduction of carbon dioxide to carbon monoxide; (b) reacting the carbon monoxide produced by the carbon dioxide reduction electrolytic cell in one or more of a plurality of intermediate reactions to produce one or more intermediate chemical substances; and (c) synthesizing a polycarbonate polymer from the one or more intermediate chemical substances or one or more derivatives of the one or more intermediate chemical substances. A method comprising the above steps. [Item 19] The method according to item 18, further comprising producing bisphenol A and transporting the bisphenol A from the bisphenol A to a polycarbonate synthesis reactor, wherein the one or more intermediate chemical substances include bisphenol A. [Item 20] The method according to item 18, wherein the step of reacting the carbon monoxide produced by the carbon dioxide reduction electrolytic cell in one or more of a plurality of intermediate reactions includes reacting carbon monoxide in a gas fermentation reaction. [Item 21] The method according to item 20, wherein the gas fermentation reaction produces a product selected from the group consisting of naphtha, phenol, and acetone. [Item 22] The method according to item 18, wherein the step of reacting the carbon monoxide produced by the carbon dioxide reduction electrolytic cell in one or more of a plurality of intermediate reactions includes reacting carbon monoxide in a Fischer-Tropsch reaction to produce naphtha. [Item 23] The method according to item 18, wherein the plurality of intermediate reactions includes a naphtha cracking reaction for producing unsaturated hydrocarbons from naphtha. [Item 24] The method according to item 23, wherein the plurality of intermediate reactions further includes a Fischer-Tropsch reaction and / or a gas fermentation reaction for producing the naphtha. [Item 25] The method according to item 23, wherein the plurality of intermediate reactions further includes a cumene process reaction for producing phenol and acetone from the unsaturated hydrocarbon. [Item 26] The method according to item 25, wherein the plurality of intermediate reactions includes a bisphenol A synthesis reaction for producing bisphenol A. [Item 27] The method according to item 18, wherein the plurality of intermediate reactions includes a phosgene generation reaction for producing phosgene from the carbon monoxide and chlorine. [Item 28] The method according to item 27, further comprising a step of generating the chlorine by electrolysis in a chlor-alkali cell. [Item 29] The method according to item 28, further comprising a step of supplying electricity to the chlor-alkali cell and the carbon dioxide reduction electrolytic cell from a common power source. [Item 30] The method according to item 28, further comprising a step of delivering oxygen generated by the carbon dioxide reduction electrolytic cell to the chlor-alkali cell. [Item 31] The method according to item 18, further comprising a step of recycling carbon dioxide from at least one of the plurality of intermediate reactions to the carbon dioxide reduction electrolytic cell. [Item 32] The method according to item 31, wherein the at least one of the plurality of intermediate reactions includes a Fischer-Tropsch reaction or a gas fermentation reaction. [Item 33] The method according to item 18, wherein the carbon dioxide reduction electrolytic cell generates hydrogen and the carbon monoxide at a molar ratio of at least about 0.5:1. [Item 34] A system for producing metal formate, comprising: (a) A carbon dioxide reduction electrolyzer having a membrane electrode assembly including one or more ion-conductive polymer layers and a cathode catalyst for promoting the chemical reduction of carbon dioxide to carbon monoxide; (b) A formate synthesis reactor configured to receive the carbon monoxide produced by the carbon dioxide reduction electrolyzer and produce metal formate; and (c) One or more units configured to separate and / or purify the metal formate produced by the formate synthesis reactor . [Item 35] The system according to item 34, wherein the formate synthesis reactor is configured to receive an alkali metal hydroxide and produce the metal formate. [Item 36] The system according to item 34, wherein the one or more units configured to separate and / or purify the metal formate have a degassing device configured to remove gas from the output of the formate synthesis reactor. [Item 37] The system according to item 34, wherein the one or more ...

Claims

[Claim 1] 1. A system for producing liquid hydrocarbons from carbon dioxide, comprising: (a) a carbon dioxide reduction electrolyzer comprising: (i) a membrane electrode assembly having one or more ion-conducting polymer layers; and (ii) a cathode having a cathode catalyst layer including a cathode catalyst for promoting the chemical reduction of carbon dioxide to carbon monoxide, wherein the cathode catalyst layer is in contact with at least one of the one or more ion-conducting polymer layers; (b) a liquid hydrocarbon synthesis reactor configured to produce a liquid hydrocarbon mixture from the carbon monoxide and hydrogen; and (c) a gas separation unit configured to: (i) condense water from the cathode outlet stream of the carbon dioxide reduction electrolyzer to produce a dehumidified outlet stream; and (ii) remove carbon dioxide from the dehumidified outlet stream to produce a separated carbon dioxide stream. (d) a carbon dioxide recycle loop configured to recycle at least a portion of the separated carbon dioxide stream to the cathode of the carbon dioxide reduction electrolyzer; Equipped with The system is configured to deliver the hydrogen and the carbon monoxide from the carbon dioxide reduction electrolyzer to the liquid hydrocarbon synthesis reactor.

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