CO2 hydrogenation in a counter-flow reactor

Abstract

Systems and methods are provided for the hydrogenation of CO2 via the reverse water gas shift reaction in a counter-flow reactor environment. The counter-flow reactor environment is suitable for carrying out endothermic reactions at high temperatures, with a reactant stream passing through the reactor in a first flow direction during a first portion of the cycle, while a combustion or heating stream passes through the reactor during a second portion of the reaction cycle from the opposite direction. This may allow for efficient heating of surfaces within the reactor to provide heat for the endothermic reverse water gas shift reaction, while reducing or minimizing incorporation of combustion products into the desired reaction products.

Description

[Technical field]

[0001] The present invention relates to a system and method for carrying out CO2 hydrogenation in a reactor utilizing alternating flow directions. [Background technology]

[0002] CO2 is a common reaction product from a variety of reaction processes involving carbon compounds. In addition to the formation of CO2 during combustion, CO2 can also be formed during other refinery or chemical manufacturing processes, such as hydrocarbon reforming. Due to the prevalence of combustion and other processes that produce CO2 as a reaction product, methods that can provide for disposal of CO2 other than emission to the atmosphere are desirable.

[0003] One reaction that can consume CO2 as a reaction input is the reverse water gas shift reaction. In the reverse water gas shift reaction, in the presence of a suitable catalyst, CO2 and H2 can be converted to CO and H2O. Since this reaction is an equilibrium process, the ability to consume CO2 depends on the reaction conditions and the relative concentrations of the reactants and products. With respect to reaction conditions, the conversion of CO2 to CO is favored with increasing temperature.

[0004] A countercurrent reactor is an example of a type of reactor that is beneficial for use in processes with cyclic reaction conditions. For example, due to the endothermic nature of the reforming reaction, there is a need to introduce additional heat to the reforming reaction environment on a consistent basis. A countercurrent reactor can provide an efficient manner to introduce heat to the reaction environment. After a portion of the reaction cycle is used for reforming or another endothermic reaction, a second portion of the reaction cycle can be used for combustion or another exothermic reaction to add heat to the reaction environment in preparation for the next reforming step. US Pat. No. 5,399,433 and US Pat. No. 5,499,433 provide examples of using countercurrent reactors to perform various endothermic processes in cyclic reaction environments.

[0005] US Pat. No. 5,399,433 describes the production of synthesis gas in a counter-flow reactor by steam reforming followed by incomplete combustion of the residual hydrocarbons at high temperature and pressure. In addition to providing additional synthesis gas, the incomplete combustion provides heat to the reactor. In the method described in US Pat. No. 5,399,433, flow reversal is achieved by alternating the ends of the reactors used for input of reactant streams for carrying out steam reforming and incomplete combustion. The resulting synthesis gas can then be used for the production of methanol.

[0006] US Pat. No. 5,399,633 describes an in-situ vaporizer and recuperator suitable for use in alternating flow systems such as pressure swing reformers. [Prior art documents] [Patent documents]

[0007] [Patent Document 1] U.S. Patent No. 7,815,873 [Patent Document 2] U.S. Pat. No. 8,754,276 [Patent Document 3] U.S. Patent No. 7,740,289 [Patent Document 4] US Patent Application Publication No. 2012 / 0111315 Summary of the Invention [Means for solving the problem]

[0008] In various embodiments, a method is provided for carrying out a reaction in a reverse flow reactor. The method includes reacting a fuel mixture comprising a fuel, 0.1% by volume or more O2, and a diluent under combustion conditions in a combustion zone in the reverse flow reactor to form a combustion gas, and heating the conversion zone to an average conversion zone temperature of 400° C. or more. The conversion zone can include a catalyst composition. The method further includes exposing a reactant stream comprising 5.0% by volume to 50% by volume CO2, 20% by volume or more H2, with a ratio of H2 to CO2 of 2.0 or more, to the catalyst composition in the conversion zone under reverse water gas shift reaction conditions to form a product stream comprising H2O and CO. The direction of flow of the reactants in the reaction zone can be reversed relative to the direction of flow of the fuel mixture. [Brief description of the drawings]

[0009] [Figure 1] 1 shows CO2 conversion during the water-gas shift portion of the reaction cycle in a counter-flow reactor under a first set of conditions. [Diagram 2] 1 shows methane formation during the water-gas shift portion of the reaction cycle in a counter-flow reactor under a first set of conditions. [Diagram 3] 13 shows CO2 conversion during the water-gas shift portion of the reaction cycle in a counter-flow reactor under a second set of conditions. [Figure 4] 13 shows CO2 conversion during the water-gas shift portion of the reaction cycle in a counter-flow reactor under an additional set of conditions. [Diagram 5] 13 shows the methane content of the reaction product during the water-gas shift portion of the reaction cycle in a counter-current reactor under an additional set of conditions. [Figure 6] An example of the operation of a counter-flow reactor is given. [Figure 7] An example of a counter-flow reactor is shown. [Figure 8] 1 shows an example of the temperature profile in a counter-flow reactor during a reaction cycle involving the reverse water gas shift reaction. DETAILED DESCRIPTION OF THE PREFERRED EMBODIMENTS

[0010] All numerical values ​​within the detailed description and claims herein are modified by the stated value being "about" or "approximately" to account for experimental error and variations that would be expected by one of ordinary skill in the art.

[0011] In various embodiments, systems and methods are provided for the hydrogenation of CO in a counter-flow reactor environment suitable for carrying out endothermic reactions at high temperatures, where a reactant stream passes through the reactor in a first part of the cycle in a first flow direction, while a combustion or heating stream passes through the reactor during a second part of the reaction cycle from the opposite direction, allowing efficient heating of surfaces within the reactor to provide heat for the endothermic reverse water gas shift reaction, while reducing or minimizing incorporation of combustion products into the desired reaction products.

[0012] The reverse water gas shift reaction is a reaction that allows the conversion of a relatively inert compound (CO2) into a compound that is susceptible to a wider range of reactions. For example, catalysts and / or reactor configurations for carrying out Fischer-Tropsch synthesis or methanol synthesis can easily use CO as a reactant, but typically cannot use CO2. The reverse water gas shift reaction converts CO2 and H2 into CO and H2O. As a result, by carrying out the reverse water gas shift reaction, it is possible to use CO2 and H2 to form synthesis gas, which can then be used in synthesis reactions such as Fischer-Tropsch synthesis, methanol synthesis, or dimethyl ether synthesis.

[0013] Performing the reverse water gas shift reaction in a counter-current reactor can provide a variety of advantages. Some advantages relate to the high temperatures achievable in a counter-current reactor. The reverse water gas shift reaction is part of the same equilibrium as the water gas shift reaction. At that equilibrium, the formation of CO and H2O is favored with increasing temperature. In particular, the equilibrium conversion of CO2 passes a minimum at about 600°C, due to competing equilibrium reactions for the formation of methane. By performing the reverse water gas shift reaction at temperatures of 700°C or higher, or 800°C or higher, or 900°C or higher (e.g., up to 1600°C, or possibly higher), the equilibrium conversion of CO2 can be increased while operating at temperatures at a relatively rapid rate. In various embodiments, the amount of CO2 in the reaction product can be from 0.5% to 5.0% by volume, or from 0.5% to 3.0% by volume, or from 0.5% to 2.5% by volume, or from 1.0% to 5.0% by volume, or from 1.0% to 3.0% by volume, or from 1.0% to 2.5% by volume.

[0014] The high temperatures achievable in the countercurrent reactor can also cause the formation of CO and H2O under reverse water gas shift conditions. This allows for the production of syngas compositions from the countercurrent reactor that are desirable for downstream synthesis processes. In various embodiments, the molar ratio of H2 to CO in the product from the countercurrent reactor process can be 1.9 to 3.1, or 2.0 to 3.0. In this discussion, this ratio can also be described as the syngas ratio. A reaction product having a syngas ratio closer to 2.0 can be beneficial, for example, for Fischer-Tropsch synthesis and methanol synthesis. In other applications, such as using the syngas as a feed for fermentation (i.e., chemical production via algae or another bioreaction route), a syngas ratio closer to 3.0 can be beneficial.

[0015] In some embodiments, another advantage of performing the reverse water gas shift reaction in a countercurrent reactor may be related to the type of catalyst. For example, it has been discovered that catalysts suitable for reforming hydrocarbons at temperatures of 600° C. or higher may also be suitable for performing the reverse water gas shift reaction. It has further been discovered that it is possible to add methane to the CO2 feed to enable reforming and reverse water gas shift feed conversion at pressures of 200 psig (about 1400 kPa-g) or higher, or 290 psig (about 2000 kPa-g) or higher.

[0016] Yet another advantage of using a counter-flow reactor to perform the reverse water gas shift reaction can be related to heat management and reaction product management. During the regeneration step of the reaction cycle in the counter-flow reactor, heat is added to the internal surfaces and / or structures within the reactor by performing combustion in the combustion zone. The internal surfaces and structures are exposed to the heat resulting from the combustion as the heat is distributed by the diluent gas flow. This provides direct internal heating that can reduce or minimize energy losses compared to indirect methods where the external parts of the reactor and / or reaction zone are heated. Heat diffusion then transports the heat from the external locations to the internal surfaces where the endothermic reaction (e.g., the reverse water gas shift reaction) is performed.

[0017] Yet another advantage of using a counter-flow reactor to perform the reverse water gas shift reaction can relate to reaction product management. In a counter-flow reactor, the regeneration step to add heat to the reactor occurs at a separate time from the reaction step to perform the reverse water gas shift reaction (or another desired endothermic reaction). As a result, combustion products from the regeneration step can be separated from the reaction products of the reverse water gas shift reaction.

[0018] When carrying out endothermic reactions at high temperatures, counter-flow reactors and / or other reactors with counter-current flows at different stages of the reaction cycle may be useful. Examples of high temperatures correspond to temperatures of 700° C. or higher, or 800° C. or higher, or 900° C. or higher, for example up to 1600° C. or possibly higher. The flow from the first direction corresponds to a combustion or regeneration flow and can be used to heat the reaction zone in the reactor to the desired temperature. The flow in the counter-current direction can then be used to pass the reagents for the desired reaction. The heat stored in the reactor during the regeneration step is used to provide heat for the desired endothermic reaction.

[0019] In this discussion, unless otherwise stated, all volume ratios correspond to volume ratios where the ratio amounts are stated based on volumes at standard temperature and pressure (20° C., 100 kPa). This allows the volume ratios to be consistently stated even though the two combustion gas volumes being compared may be at different temperatures and pressures. When volume ratios are stated with respect to the combustion gases fed into the reactor, the equivalent flow rate of the gases during unit time under standard conditions can be used for comparison.

[0020] In this discussion, unless otherwise specified, the temperatures of the reactor surface and / or within the reactor described herein correspond to the temperature at the peak temperature location of the conversion zone after the end of the regeneration step. The peak temperature of the conversion zone after the end of the regeneration step can be the highest temperature to which the reactor surface is exposed during the reaction cycle. (Depending on the location where the fuel is first burned during the regeneration step, the absolute peak temperature of the reactor can be outside the conversion zone.) The location of the conversion zone where the peak temperature occurs after the end of the regeneration step can then be used as a reference location to characterize the reactor temperature at other times during the reaction cycle. Note that the temperature of the combustion gases leaving the reactor during regeneration is different from the temperature at the peak temperature location.

[0021] Process Example - Reverse Water Gas Shift and Regeneration An example of a reaction that can be carried out in a reverse flow reactor system is the reverse water gas shift reaction. As a general overview of the operation of a swing reactor such as a reverse flow reactor, the regeneration step or portion of the reaction cycle can be used to provide heat to the reactor. The reverse water gas shift reaction can then occur in the reactor during the conversion step or portion of the cycle, and the reverse water gas shift reaction consumes the heat provided during the reactor regeneration step. During reactor regeneration, fuel, oxidant and diluent are introduced into the reactor from the regeneration end of the reactor. The beds and / or monoliths of the regeneration portion of the reactor can absorb heat, but typically do not contain a catalyst for the reverse water gas shift reaction, such as a reforming catalyst. As the fuel, oxidant and diluent pass through the regeneration portion, heat is transferred from the regeneration portion to the fuel and oxidant. Combustion does not occur instantly, but instead the location of combustion is controlled to occur in the middle portion of the reactor. During the regeneration step, the flow of fuel, oxidant and diluent continues, and additional transfer of heat resulting from combustion to the reforming end of the reactor is conducted.

[0022] After a sufficient time, the combustion reaction is stopped. Any remaining combustion products and / or reactants can be optionally removed. The conversion step or portion of the reaction cycle (i.e., the reverse water gas shift portion) can then begin. The reactants for conversion can be introduced at the conversion end of the reactor, and thus can effectively flow in a counter-direction to the flow during regeneration. The bed and / or monolith in the conversion portion of the reactor can include a catalyst suitable for performing a reverse water gas shift reaction. In various aspects, at least a portion of the catalyst can correspond to a catalyst formed from a ceramic composition, as described herein. As the conversion occurs, the heat introduced into the conversion zone during combustion can be consumed by the endothermic reverse water gas shift reaction. Optionally, if a reforming catalyst is used, some heat can also be consumed by reforming of the hydrocarbons present in the reactants for the conversion reaction. After leaving the conversion zone, the conversion products (and unreacted reactants) are no longer exposed to the conversion catalyst. As the conversion products pass through the regeneration zone, heat can be transferred from the products to the regeneration zone. After a sufficient time, the conversion process can be stopped, the remaining conversion products can optionally be recovered or purged from the reactor, and the cycle can begin again with the regeneration step.

[0023] The conversion reaction carried out in the reactor can correspond to the reverse water gas shift reaction (RWGS) shown in equation (1). Optionally, some reforming may also occur if a reforming catalyst is used to catalyze the reverse water gas shift reaction. Equation (2) shows an example of a stoichiometry for steam reforming. Note that equations (1) and (2) correspond to equilibrium processes. (1)RWGS:H2+CO2<=>CO+H2O (2) Steam reforming: CH4 + H2O = CO + 3H2

[0024] It should be noted that the ratio of H2 to CO in the syngas produced by equation (1) depends on the temperature and the amount of CO2 and H2O present in the reaction environment. If reforming also occurs, the equilibrium for the reforming reaction may also affect the resulting syngas ratio.

[0025] Most reforming catalysts, such as rhodium and / or nickel, can function as water-gas shift catalysts. Thus, if the reaction environment for producing H2 and CO also includes HO and / or CO2, the initial stoichiometry from the reforming reaction may be altered based on the water-gas shift equilibrium. This equilibrium is also temperature dependent, with higher temperatures being favored for the production of CO and HO. However, higher temperatures can improve the rate at which equilibrium is reached.

[0026] One of the challenges of running an endothermic reaction on a commercial scale can be providing heat to the reaction in an efficient manner while reducing or minimizing the introduction of additional components to the desired product. In a cyclic reaction system, such as a counter-flow reactor system, heat can be provided in a desired manner by having a cycle that includes a conversion step and a regeneration step. During the regeneration step, combustion can be performed in a selected region of the reactor. A gas flow during regeneration (such as a gas flow with diluent) can assist in the transfer of heat from the combustion region to additional parts of the conversion region in the reactor. The conversion step in the cycle can be a separate step, so that the incorporation of products from combustion and / or products from conversion into the reactants can be reduced or minimized. The conversion step can consume heat, thereby reducing the temperature of the conversion region. Once the products from the conversion pass through the reactor, the conversion products can pass through a second region lacking reforming or water gas shift catalyst. This allows the reaction products to be cooled before exiting the reactor. The heat transferred from the conversion products to the reactor can then be used to increase the temperature of the reactants for the next combustion or regeneration step.

[0027] In various embodiments, the feed for the reverse water gas shift reaction can include 5.0% to 50% CO2 by volume and 20% to 80% H2 by volume. The amount of H2 in the feed can be sufficient to provide a ratio of H2 to CO2 in the feed of 2.0 or more, or 3.0 or more, such as up to 5.0, or possibly more. Depending on the embodiment, 5.0% to 40% H2O by volume can be included in the feed. In some embodiments, 0.5% to 20% reformable hydrocarbons, such as 1.0% to 20% methane by volume, can be included in the feed. The above values ​​correspond to the volume percentages of the reactants, excluding any diluents that may be present. For example, when determining the volume percentages of the reactants in the feed, any N2 present in the feed can be excluded.

[0028] Within the conversion zone of a counter-flow reactor, the temperature can vary across the zone due to the nature of how heat is added to the reactor and / or due to the kinetics of the conversion reaction. The hottest part of the zone is typically found near the middle of the reactor. This middle part can be described as the mixing zone where combustion starts during regeneration. If the monolith with the conversion catalyst extends into the mixing zone, at least a portion of the mixing zone can represent a portion of the conversion zone. As a result, the location where combustion starts during regeneration can typically be near the end of the conversion zone in the reactor. Moving from the center of the reactor to the end of the reactor, the temperature can decrease. As a result, the temperature at the beginning of the conversion zone (at the end of the reactor) can be lower than the temperature at the end of the conversion zone (at the middle of the reactor).

[0029] As the conversion reaction occurs, the temperature within the conversion zone can decrease. The rate at which the temperature decreases can be related to the kinetics of the amount of CO2 available for conversion and / or the temperature at a given location within the conversion zone. As the feed moves through the conversion zone, reactants in the feed can be consumed, thereby decreasing the amount of conversion occurring at downstream locations. However, an increase in the temperature of the conversion zone as reactants move across the conversion zone can lead to an increase in the reaction rate.

[0030] The conversion step in the reaction cycle can be stopped for regeneration when a sufficient portion of the conversion zone has a temperature below 800°C, or below 700°C, or below 600°C (e.g., up to 400°C, or even lower, if possible). For example, the conversion step can be stopped when the average temperature of the conversion zone is below 600°C, or below 500°C, or below 400°C, e.g., up to 300°C. Alternatively, the conversion portion of the reaction cycle can be stopped based on the amount of reaction time such that the amount of heat consumed during conversion (and heat lost to the environment) is approximately balanced with the amount of heat added during regeneration, based on the amount of heat introduced into the reactor during regeneration. After the conversion process is stopped, any remaining synthesis gas product still in the reactor can be optionally recovered before starting the regeneration step of the reaction cycle.

[0031] The regeneration process can then begin. During regeneration, fuels such as methane, natural gas, or H2 and oxygen can be introduced into the reactor and burned. In some embodiments, a diluent can also be introduced to facilitate the transfer of heat within the reactor. The location at which the fuel and oxidant can be mixed can be controlled by any convenient method, such as by introducing the fuel and oxidant through separate channels. During regeneration, the non-converting end of the reactor can be kept at a lower temperature by delaying combustion until the reactants reach the center of the reactor. This can result in a temperature peak in the middle of the reactor. The temperature peak can be within a portion of the reactor that also contains the reforming catalyst. During the regeneration cycle, the temperature in the reactor can be increased sufficiently to allow for conversion reactions during the conversion portion of the cycle. This can result in a peak temperature in the reactor of 1100° C. or higher, or 1200° C. or higher, or 1300° C. or higher, or potentially even higher.

[0032] The relative lengths of time and reactant flow rates for the conversion and regeneration portions of the process cycle can be selected to provide a balance between the heat provided during regeneration and the heat consumed during conversion. For example, one option can be to select a conversion step having a similar length as the regeneration step. Based on the flow rates of H2 and / or CO2 during the reforming step, the endothermic heat demand for the conversion reaction can be determined. This heat demand can then be used to calculate the flow rates for the combustion reactants during the regeneration step. Of course, in other embodiments, the heat balance between conversion and regeneration can be determined in other ways, such as by determining the desired flow rates of the reactants and then selecting the cycle length to provide a balance between the heat provided by regeneration and the heat consumed during conversion.

[0033] In addition to providing heat, the reactor regeneration step during the reaction cycle can also allow for coke removal from the catalyst in the conversion zone. In various embodiments, one or more types of catalyst regeneration can potentially occur during the regeneration step. One type of catalyst regeneration can correspond to the removal of coke from the catalyst. During reforming, a portion of the CO2 introduced into the conversion zone can form coke instead of forming CO. This coke can potentially block access to catalytic sites (such as metal sites) of the catalyst. In some embodiments, the formation rate can be increased in parts of the conversion zone exposed to higher temperatures, such as parts of the reforming zone exposed to temperatures of 800°C or higher, or 900°C or higher, or 1000°C or higher. During the regeneration step, oxygen can be present when the temperature of the reforming zone is increased. At the temperatures reached during regeneration, at least a portion of the coke generated during conversion can be removed as CO or CO2.

[0034] Due to temperature variations within the reactor, several options can be used to characterize the temperature within the reactor and / or within the conversion zone of the reactor. One option for characterizing the temperature can be based on the average bed or average monolith temperature within the reforming zone. In a practical setting, the determination of the temperature within the reactor requires the presence of a measuring device such as a thermocouple. Rather than attempting to measure the temperature within the conversion zone, the average (bed or monolith) temperature within the conversion zone can be defined based on the average of the temperature at the beginning of the conversion zone and the temperature at the end of the conversion zone. Another option can be to characterize the peak temperature within the conversion zone after the regeneration step of the reaction cycle. In general, the peak temperature can occur at or near the end of the conversion zone and may depend on where combustion begins in the reactor. Yet another option can be to characterize the temperature difference at a given location within the reaction zone at different times within the reaction cycle. For example, the temperature difference can be determined between the temperature at the end of the regeneration step and the temperature at the end of the conversion step. Such temperature differentials may be characterized at the location of the peak temperature within the reactor, at the entrance to the conversion region, at the exit from the conversion region, or at any other convenient location.

[0035] In various embodiments, the reaction conditions for carrying out the reverse water gas shift reaction can include one or more of: an average conversion zone temperature in the range of 400° C. to 1200° C. (or higher); a peak temperature in the conversion zone of 800° C. to 1500° C.; a temperature difference at the location of the peak temperature between the end of the regeneration step and the end of the subsequent conversion step of 25° C. or higher, or 50° C. or higher, or 100° C. or higher, or 200° C. or higher, e.g., up to 800° C., or possibly even higher; a temperature difference at the inlet to the conversion zone between the end of the regeneration step and the end of the subsequent conversion step of 25° C. or higher, or 50° C. or higher, or 100° C. or higher, or 200° C. or higher, e.g., up to 800° C., or possibly even higher; and / or a temperature difference at the outlet from the conversion zone between the end of the regeneration step and the end of the subsequent conversion step of 25° C. or higher, or 50° C. or higher, or 100° C. or higher, or 200° C. or higher, e.g., up to 800° C., or possibly even higher.

[0036] With respect to the average conversion zone temperature, in various embodiments, the average temperature of the conversion zone at the start of the conversion step can be 400° C. to 1500° C., or 400° C. to 1200° C., or 500° C. to 1500° C., or 500° C. to 1200° C., or 600° C. to 1500° C., or 600° C. to 1200° C., or 900° C. to 1500° C., or 900° C. to 1200° C., or 1000° C. to 1500° C., or 1000° C. to 1200° C. Additionally or alternatively, with respect to the peak temperature of the conversion zone (possibly corresponding to a position of the conversion zone near a position for combustion of the regenerated reactants), the peak temperature can be 800° C. to 1500° C., or 800° C. to 1400° C., or 1200° C. to 1500° C., or 1200° C. to 1400° C.

[0037] Additionally or alternatively, the reaction conditions for the reverse water gas shift reaction are a pressure of 0 psig to 1500 psig (10.3 MPa-g), or 0 psig to 1000 psig (6.9 MPa-g), or 200 psig (1.4 MPa-g) to 1500 psig (10.3 MPa-g), or 200 psig (1.4 MPa-g) to 1000 psig (6.9 MPa), or 290 psig (2.0 MPa-g) to 1500 psig (10.3 MPa), or 290 psig (2.0 MPa-g) to 1000 psig (6.9 MPa); and -1 ~50,000 hours -1 or 10,000 hours -1 ~50,000 hours -1 The gas space velocity of the reactants can include the gas space velocity of the reactants relative to the volume of the monolith per unit time. The volume of the monolith is defined as the volume of the monolith as if it were a solid cylinder.

[0038] Examples of counter-flow reactor configurations For endothermic reactions carried out at high temperatures, such as the reverse water gas shift reaction or hydrocarbon reforming, a counter-flow reactor can provide a suitable reaction environment to provide heat to the endothermic reaction.

[0039] In a counter-flow reactor, the heat required for the endothermic reaction may be provided by creating a high temperature thermal bubble in the center of the reactor. A two-step process can then be used in which heat is (a) added to the reactor bed or monolith via in-situ combustion and then (b) removed in-situ via an endothermic process such as reverse water gas shift, reforming, pyrolysis or steam cracking. This type of arrangement can provide the ability to consistently manage and confine the high temperature bubble in a reactor region that can withstand such conditions for extended periods of time. The counter-flow reactor system allows the first endothermic and regenerative process to be carried out in a substantially continuous manner.

[0040] The countercurrent reactor system may include a first and a second reactor arranged in series with respect to a common flow path and, optionally, but preferably along a common axis. The common axis may be horizontal or vertical. During the regeneration step, the reactants (e.g., fuel and oxygen, and optional diluent) may be combined or mixed in the reaction zone for in-situ combustion therein and to create a high temperature region or thermal bubble in the central portion of the reaction system. The thermal bubble may correspond to a temperature that is at least about the initial temperature of the endothermic reaction. Typically, the temperature of the thermal bubble may be higher than the initial temperature for the endothermic reaction, since the temperature decreases as heat is transferred from the thermal bubble in the central portion of the reactor toward the end of the reactor. In some embodiments, the combining may be enhanced by a reactant mixer, optionally located between the first and second reactors, that mixes the reactants to promote substantially complete combustion / reaction at the desired location. The combustion process can be carried out for a long enough period of time so that the flow of the first and second reactants in the first reactor displaces a substantial portion (if desired) of the heat (e.g., heat bubbles) generated by the reaction into and at least partially, but preferably not entirely, through the second reactor, also serving to avoid waste heat and heating of the second reactor. The combustion gases may be exhausted through the second reactor, but preferably most of the heat is retained within the second reactor. The amount of heat displaced into the second reactor during the regeneration step can also be limited or determined by the desired exposure time or space velocity that the feed gas has in the endothermic reaction environment.

[0041] After regenerating or heating the second reaction medium (which may include and / or correspond to the ceramic catalyst composition described herein), in the next / reverse step or cycle, the reactants for the endothermic reaction may be fed into the second reactor or may flow from a direction opposite to the flow direction during the heating step. For example, in a reforming process, methane (and / or natural gas and / or another hydrocarbon) may be fed into or flow into the second reactor. The methane may be contacted with the heated second reactor and mixer medium in the thermal bubble region, transferring heat for reaction energy to the methane.

[0042] For some embodiments, the basic two-step asymmetric cycle of a countercurrent regenerative bed reactor system is shown in Figures 6A and 6B of Figure 6 for a reactor system having two zones / reactors; a first or recuperator / quench zone (7) and a second or reaction zone (1). Both the reaction zone (1) and the recuperator zone (7) can contain a regenerative monolith and / or other regenerative structure formed from a doped ceramic composition. Regenerative monolith or other regenerative structure, as used herein, includes materials that are effective in storing and transferring heat and are effective for carrying out chemical reactions. The regenerative monolith and / or other structure can correspond to any convenient type of material suitable for storing heat, transferring heat, and catalyzing reactions. Exemplary structures may include layered or packed ceramic beads or spheres, ceramic honeycomb materials, ceramic tubes, extruded monoliths, etc., provided they are capable of withstanding prolonged exposure to temperatures above 1200° C., or above 1400° C., or above 1600° C. while maintaining integrity, functionality, and allowing some operating margin. In some embodiments, catalytic ceramic monoliths and / or other catalytic ceramic structures may be used without the presence of an additional washcoat.

[0043] To facilitate the explanation of Figure 6, a reactor for a general endothermic reaction is described herein, it being understood that a variety of convenient types of endothermic reactions can be generally carried out using a counter-flow reactor such as the reactor shown in Figure 6.

[0044] As shown in FIG. 6B of FIG. 6, at the start of the "react" step of the cycle, the second end 5 of reaction zone 1 (also referred to herein as the second reactor) can be at a higher temperature compared to the first end 3 of reaction zone 1, and at least a portion of the recuperator or quench zone 7 (also referred to herein as the first reactor) (including first end 9) can be at a lower temperature than reaction zone 1 to provide a quenching effect to the resulting product. In embodiments where the reactor is used to perform countercurrent reforming, reactant feed is introduced to the first end 3 of conversion or reaction zone 1 via conduit 15.

[0045] The feed stream from inlet 15 can absorb heat from reaction zone 1 and endothermically react to produce the desired conversion product. As this step progresses, a shift in temperature profile 2 can be created based on the heat transfer characteristics of the system, as indicated by the arrows. If the ceramic catalyst monolith / other catalyst structure is designed with proper heat transfer capabilities, this profile can have a relatively sharp temperature gradient, and such a gradient can move across reaction zone 1 as the conversion step progresses. In some embodiments, a sharper temperature gradient profile results in improved control of reaction conditions.

[0046] The effluent from the conversion reaction, which may include unreacted feed components as well as reaction products, may exit the reaction zone 1 through second end 5 at an elevated temperature, pass through the recuperator reactor 7, and pass through second end 11 and exit first end 9. The recuperator 7 may initially be at a lower temperature than the reaction zone 1. As the products from the reforming reaction (and optionally unreacted feed) pass through the recuperator zone 7, the gas may be quenched or cooled at first end 9 to a temperature substantially close to the temperature of the recuperator zone, which in some embodiments may be about the same temperature as the regenerator feed introduced through conduit 19 into the recuperator 7 during the second step of the cycle. As the reformer effluent cools in the recuperator zone 7, a temperature gradient 4 may be created in the regenerator layer of the zone and may travel throughout the recuperator zone 7 during this step. The quench can heat the recuperator 7, which can be cooled again in a second step to provide for another quench run and prevent the size and location of the thermal bubble from increasing progressively through the quench reactor 7. After the quench, the reaction gas can exit the recuperator at 9 through conduit 17 and can be processed for separation and recovery of various components.

[0047] The second step of the cycle, called the regeneration step, can then begin with the reintroduction of the first and second regeneration reactants through conduit 19. The first and second reactants can be passed separately through the heating recuperator 7 towards the second end 11 of the recuperator 7, where they can be combined for exothermic reaction or combustion at or near the central region 13 of the reactor system.

[0048] An example of the regeneration step is shown in FIG. 6B of FIG. 6. Regeneration can involve the transfer of substantial recovered heat from the recuperator zone 7 to the reaction zone 1 to thermally regenerate the reaction bed 1 for a subsequent reaction cycle. Regeneration gas / reactants can enter the recuperator zone 7, for example, via conduit 19, pass through the recuperator zone 7, and flow to the reaction zone 1. In so doing, temperature gradients 6 and 8 can travel throughout the bed, as shown by the arrows on the exemplary graph in FIG. 6B, and are similar, but in the opposite direction, to the graph of the temperature gradients developed during a reaction cycle in FIG. 6B of FIG. 6. Fuel and oxidant reactants can be combusted in a region near the interface 13 of the recuperator zone 7 and the reaction zone 1. The heat recovered from the recuperator zone along with the heat of combustion can be transferred to the reaction zone and thermally regenerate the regenerated reaction monolith and / or bed 1 disposed therein.

[0049] In some embodiments, some conduits within the channel may carry a mixture of the first and second reactants, at least in part due to some mixing at the first end (17) of the first reactor. However, the number of conduits carrying a combustible mixture of the first and second reactants can be low enough so that the majority of the stoichiometrically reactable reactants do not react until they leave the second end of the first reactor. The axial location of the onset of combustion or exothermic reactions carrying the mixture of reactants in those conduits can be controlled by a combination of temperature, time, and fluid dynamics. Fuel and oxygen typically require temperature- and mixture-dependent autoignition times to combust. However, some reactions may occur within the axial portion of the conduits carrying the mixture of reactants. However, this reaction is acceptable because the number of channels with such reactions is small enough that there is only an acceptable or insignificant level of impact on the overall heat balance within the reactor. The design details of a particular reactor can be selected to avoid mixing of the reactants in as many conduits as reasonably possible.

[0050] FIG. 7 shows another exemplary reactor system that may be suitable in some applications for controlling and postponing the combustion of fuel and oxidant to achieve efficient regenerative heat. FIG. 7 shows a single reactor system operating in a regenerative cycle. The reactor system may be considered as including two reactor regions. Recuperator 27 can be the region where quenching primarily occurs and provides a substantially isolated flow path or channel for both quench reactant gases to travel through the reactor medium without incurring combustion until the gases arrive near or within the reactor core of FIG. 6. Conversion reactor 2 can be the reactor where regenerative heating and reverse water gas shift (or another endothermic reaction) primarily occur and may be considered as the second reactor for purposes of this specification. Although the first and second reactors of the reactor system are identified as separate and distinct reactors, it is understood that the first and second reactors may be fabricated, provided, or otherwise combined into a common single reactor layer, whereby the reactor system may be described as including only a single reactor integrating both cycles within the reactor. The "first reactor" and "second reactor" may simply refer to the respective regions within the reactor system whereby the respective regeneration, conversion, quench, etc. steps occur and do not require that separate components be utilized for the two reactors. However, various embodiments may include reactor systems whereby the recuperator reactor includes the conduits and channels described herein and the conversion reactor may similarly possess conduits. Additionally or alternatively, some embodiments may include a conversion reactor layer that is disposed separately from the recuperator reactor layer and may even include different materials than the recuperator reactor layer.

[0051] As discussed above, the first reactor or recuperator 27 may include various gas conduits 28 for separately flowing two or more gases after entering the first end 29 of the recuperator 27 and through a regenerator bed disposed therein. The first gas 30 may enter the first end of the plurality of flow conduits 28. In addition to providing flow channels, the conduits 28 may also include flow barriers (e.g., effectively functioning as conduit walls) effective to prevent cross-flow and mixing between the first and second reactants and to keep the primary reactants effectively separated from one another until mixing is permitted. As discussed above, each of the first and second channels may include multiple channels or flow paths. The first reactor may also include multiple substantially parallel flow segments, each of which includes separate first and second channels.

[0052] In some embodiments, the recuperator may be comprised of one or more extruded honeycomb monoliths, as described above. Each monolith may provide a flow channel (e.g., a flow path) for one of the first or second reactants. Each channel preferably includes a plurality of conduits. Alternatively, the monolith may include one or more channels for each reactant, with one or more groups of channels or conduits for flowing one or more streams of the reactants, while the remainder of the conduits for flowing one or more streams of the other reactants. It is recognized that at the interface between the channels, multiple conduits may carry a mixture of the first and second reactants, but this number of conduits is proportionately small.

[0053] In embodiments in which a monolith is used, the monolith may have any convenient shape suitable for use as a catalytic surface. One example of a monolith may be an extruded honeycomb monolith. A honeycomb monolith may be an extruded structure that includes many (e.g., a plurality, meaning two or more) small gas passages or conduits arranged in parallel in a manner with thin walls between them. Small reactors may include a single monolith, but larger reactors can include multiple monoliths, and even larger reactors may be substantially filled with an array of multiple honeycomb monoliths. Each monolith may be formed by extruding a monolith block with a shaped (e.g., square or hexagonal) cross section, and stacking such blocks on top, back, and next to each other in two or three dimensions. Monoliths may be attractive as reactor internals because they offer high heat transfer performance with minimal pressure drop.

[0054] In some embodiments, the honeycomb monolith can be characterized as having an open frontal area (or geometric void volume) of 25% to 80%, or 25% to 55%, and a conduit density of 50 to 2000 pores or cells per square inch (CPSI), or 100 to 900 cells per square inch, or 100 cells per square inch to 600 cells per square inch. For example, in one embodiment, the conduits may have a diameter / characteristic cell lateral length of only a few millimeters, on the order of 1 millimeter. The reactor media component, such as the monolith or another layer media, may be 50 ft2 or 100 ft2 or 100 ft3 or 100 ft4 or 100 ft5 or 100 ft6 or 100 ft2 or 100 ft3 or 100 ft4 ... -1 ~3000 feet -1 (Approx. 0.16km -1 ~about 10km -1 ) or 100 feet - 1 to 2500 feet -1 (about 0.32km -1 ~about 8.2km -1 ), or 200 feet -1 ~2000 feet -1 (Approx. 0.65km -1 ~about 6.5km -1) range of average wetted surface area per unit volume. These relatively high surface area per unit volume values ​​can help achieve relatively fast changes in temperature in the reactor, as generally indicated by the relatively steep slopes in the exemplary temperature gradient profile graphs shown in Figures 12(a) or 12(b) of Figure 6.

[0055] The reactor medium components have a high volumetric heat transfer coefficient (e.g., 0.02 cal / cm 3 sec °C or more, or 0.05 calorie / cm 3 sec °C or more, or 0.10 calorie / cm 3 sec °C or more, e.g. 0.50 cal / cm 3 sec °C, or possibly higher), have low flow resistance (low pressure drop); have an operating temperature range consistent with the maximum temperature encountered during regeneration; have high thermal shock resistance; and / or have a high volumetric heat capacity (e.g., 0.10 cal / cm 3 °C or above 0.20 cal / cm 3 °C or higher, e.g. 0.80 cal / cm 3 It is also possible to provide channels having a temperature (up to 1000 °C, or possibly higher) between the reactor and the reactor. As with the high surface area values, these relatively high volumetric heat transfer coefficient values ​​and / or other properties can aid in achieving relatively fast changes in temperature in the reactor, as generally shown, for example, by the relatively steep slopes in the exemplary temperature gradient profile graphs in Figures 12(a) and 12(b) of Figure 6. The values ​​listed are averages based on the volume of the reactor used for transport of the reactants.

[0056] In various embodiments, an adequate heat transfer rate can be characterized by a heat transfer parameter (ΔTHT) below 500° C., or below 100° C., or below 50° C. The parameter ΔTHT, as used herein, is the ratio of the bed average volumetric heat transfer rate required for recuperation to the volumetric heat transfer coefficient (hv) of the bed. The volumetric heat transfer rate (e.g., calories / cm) that is sufficient for recuperation is defined as the ratio of the bed average volumetric heat transfer rate (hv) to the volumetric heat transfer coefficient (hv) of the bed. 3The gas flow rate (e.g., g / sec) can be calculated as the product of the gas heat capacity (e.g., calories / g °C) and the desired end-to-end temperature change (excluding any reaction, e.g., °C), and then multiplying this amount by the volume (e.g., cm) of the reactor (or portion of the reactor) through which the gas is moving. 3 The volumetric heat transfer coefficient (hv) of the layer can be calculated by dividing it by an area-based factor (e.g., calories / cm 2 sec °C) and the surface area for heat transfer (av, e.g. cm 2 / cm 3 )

[0057] In some embodiments, a washcoat can be added to the formed, sintered ceramic composition. The washcoat allows the sintered ceramic composition to be impregnated with additional catalytic metals.

[0058] One option for incorporating additional catalytic metals into the washcoat can be to impregnate the catalytic support with the additional catalytic metal, for example by incipient wetness impregnation. Impregnation can be carried out in an aqueous solution of a suitable metal salt or other catalytic metal precursor, such as platinum tetramine nitrate or rhodium nitrate hydrate. The impregnated support can then be dried and / or calcined for decomposition of the catalytic metal precursor. Various temperature profiles can potentially be used for the heating step. To dry the support, one or more initial drying steps can be used, such as heating at a temperature between 100° C. and 200° C. for 0.5 hours to 24 hours. Depending on the nature of the impregnated catalytic metal compound, the calcination to decompose the catalytic metal precursor compound can be at a temperature between 200° C. and 800° C. for 0.5 hours to 24 hours. Depending on the precursor of the catalytic metal, the drying step and / or the decomposition calcination step can be optional. Examples of additional catalytic metals include, but are not limited to, Ni, Co, Fe, Pd, Rh, Ru, Pt, Ir, Cu, Ag, Au, Zr, Cr, Ti, V, W, Mo, Nb, Mn, Sr, La, and combinations thereof.

[0059] In another embodiment, reaction media other than monoliths may be used whereby the channel conduit / flow path may include, but is not limited to, labyrinthine, multi-flow paths having channels through a portion of the reactor, conduits, tubes, slots and / or pore structures, and may include, for example, barrier portions along the outer surface of the segments or within the subsegments that are substantially not effectively permeable to gases and / or other means suitable for preventing cross-flow between the reactant gases and for maintaining the first and second reactant gases substantially separated from each other while passing axially through the recuperator 27. For such embodiments, complex flow paths may result in an extended effective flow path, increased surface area, improved heat transfer. Such designs may be preferred for reactor embodiments having a relatively short axial length through the reactor. A longer axial length of the reactor may experience an increased pressure drop across the reactor. However, for such embodiments, the porous and / or permeable media may include, for example, at least one of a packed bed, an array of tiles, a permeable solid media, a substantially honeycomb-type structure, a fiber array, and a mesh-type lattice structure.

[0060] In some embodiments, the countercurrent reactor can include some type of device or method to direct one flow of reactants to a selected portion of the conduit. In the exemplary embodiment of FIG. 7, the gas distributor 31 can direct the second gas stream to a second gas flow channel, shown herein as channel 33, that is substantially isolated from or not in fluid communication with the first gas channel. As a result, at least a portion of the gas stream 33 can remain isolated from the gas stream 30 during the axial movement of the recuperator 27. In some embodiments, the regenerator layer and / or monolith in the recuperator region can include channels with gas or fluid barriers that isolate the first reactant from the second reactant. Thereby, both of the at least two reactant gases traveling through the channel means can pass completely through the regenerator layer, quenching the regenerator layer and allowing the reactant gases to absorb heat and then combine to react with each other in the combustion region.

[0061] In various embodiments, the gases 30 and 32 (including fluids) can each contain a component that reacts with the components of the other reactants 30 and 32, resulting in an exothermic reaction when combined. For example, each of the first and second reactants can contain one fuel gas and one oxidant gas that combusts when combined with the other of the fuel and oxidant. By keeping the reactants substantially separated, the location of the heat generated by the exothermic reaction can be controlled. In some embodiments, "substantially separated" can be defined as meaning that at least 50 percent, or at least 75 percent, or at least 90 percent of the reactants having a minimum or limiting stoichiometrically reactable amount of reactant between the first and second reactant streams are not consumed by reaction until the time that these gases complete their axial movement in the recuperator 27. In this way, the majority of the first reactant 30 can be kept isolated from the majority of the second reactant 32, and the majority of the heat generated by the reaction of combining the reactants 30 and 32 can occur after the reactants begin to exit the recuperator 27. The reactants can be gases, although optionally some reactants may include liquids, mixtures or gas phases.

[0062] The reaction rates for these regenerative streams refer to the rate of reaction that is possible based on the overall feed stoichiometry. For example, if gas 30 contains 100 volumes of air (80 volumes of N2 and 20 volumes of O2) and gas 32 contains 10 volumes of hydrogen, the maximum stoichiometric reaction is the combustion of 10 volumes of hydrogen (H2) with 5 volumes of oxygen (O2), producing 10 volumes of H2O. In this case, if 10 volumes of hydrogen are actually combusted in the recuperator area (27), this represents 100% reaction of the regenerative stream. This is despite the presence of residual unreacted oxygen, because in this example, the unreacted oxygen is present in an amount greater than the stoichiometric requirement. Thus, in this example, hydrogen is the stoichiometrically limiting component. Using this definition, reaction of less than 50% of the regenerative stream, or reaction of less than 25%, or reaction of less than 10% may occur during the axial movement of the recuperator (27).

[0063] In various embodiments, channels 28 and 33 can comprise ceramic (including zirconium oxide), alumina or other refractory materials capable of withstanding temperatures in excess of 1200° C., or 1400° C., or 1600° C. Additionally or alternatively, channels 28 and 33 can be 50 ft. -1 ~3000 feet -1 , or 100 feet -1 ~2500 feet -1 , or 200 feet -1 ~2000 feet -1 The wetted area may be 0.01 mm or less.

[0064] Referring again briefly to FIG. 6, the reactor system can include a first reactor 7 containing a first end 9 and a second end 11, and a second reactor 1 containing a first end 3 and a second end 5. The embodiments shown in FIGS. 1 and 2 are simple illustrations provided solely for illustrative purposes and are not intended to represent comprehensive embodiments. The description of the "ends" of the reactor merely refers to the terminal portions of the reactor with respect to the axial midpoint of the reactor. Thus, the description of a gas entering or exiting the "ends" of the reactor, e.g., end 9, simply means that the gas enters or exits substantially at any of a variety of points along the axis between the respective end surface of the reactor and the midpoint of the reactor, but preferably closer to the end surface than the midpoint. Thereby, one or both of the first and second reactant gases can enter at the respective end surface, while the other is supplied to the reactor at its respective end through a slot or port in the circumferential or peripheral outer surface on the respective end of the reactor. EXAMPLES

[0065] A pilot-scale reactor (approximately 12 inches / approximately 30 cm long) was used to investigate the performance of the reverse water gas shift reaction in a reverse flow reactor system containing a reforming catalyst in the conversion zone. FIG. 8 shows an example of the temperature profile in the reactor at the end of the regeneration step (triangles) and at the end of the reverse water gas shift reaction (circles). While the examples presented herein correspond to results from a single reactor, one skilled in the art will readily appreciate that the results below apply to a reaction system containing multiple reverse flow reactors. The following examples provide representative data for each operating condition. The operating conditions shown in each example were maintained for at least one hour to ensure that the results could be maintained over multiple operating cycles.

[0066] FIG. 1 shows the results from the first operating regime for the reverse water gas shift reaction. In the first operating regime, the temperature profile was similar to FIG. 8. The conversion reaction was run at 150 psig (about 1.0 MPa-g) and for about 13,300 hours using a feed containing 25% by volume CO2 and 75% by volume H2. -1 The reaction was carried out at a gas hourly space velocity (GHSV) of 1000 s. As shown in Figure 1, more than 90% of the CO2 was converted at cycle times of 15 seconds or less. Note that the decrease in CO2 conversion at times higher than 15 seconds is due to the decrease in temperature in the reactor towards the end of the reaction cycle. Therefore, by choosing a shorter time for the reaction step, the conversion could be maintained above 90%.

[0067] The effluent from the first operating condition included H2O, CO, and CH4 as reaction products. CH4 was produced by methanation or reverse reforming. Figure 2 shows the amount of methane in the reaction effluent during the reaction cycle. As shown in Figure 2, some methane production occurred, but the amount of methane production was relatively low due to the high temperature of the reverse flow reactor. In the portion of the cycle after 15 seconds, the amount of methane dropped off due to lack of CO2 conversion at the lower temperature.

[0068] Methane production in the first operating condition was relatively low, but at higher space velocities, the amount of methane production increased. To lower the amount of methane production, methane was added to the feed in the second operating condition. The second operating condition included a pressure of about 1.0 MPa-g, but the space velocity was about 2.5 times greater than the first operating condition (about 37,500 h on a total feed basis). -1 GHSV of 1000 sq. m. The feed contained approximately 20 vol.% CO2, 47 vol.% H2, 7 vol.% CH4, and 27 vol.% H2O. Figure 3 shows the CO2 conversion versus cycle time within the reaction step. As shown in Figure 3, the conversion remained near 70% throughout the cycle, with a moderate decrease as the reaction proceeded. The increase in space velocity at the second operating state is believed to be responsible for the decrease in conversion within the cycle and the corresponding increase in methane production.

[0069] FIG. 4 shows the CO2 conversion in the third and fourth operating conditions. Both operating conditions were run for 30,000 hours. -1 The third operating condition included a higher GHSV and a feed containing 15.5 vol.% CO2, 46.5 vol.% H2, 34 vol.% H2O, and 4.0 vol.% CH4. The third operating condition included a pressure of 1.0 MPa-g and the fourth operating condition had a pressure of 2.0 MPa-g.

[0070] As shown in Figure 4, varying the pressure did not have a substantial effect on the amount of CO2 conversion. However, the benefit of increasing the pressure is shown in Figure 5, which shows the amount of CH4 in the reaction effluent. As shown in Figure 5, increasing the pressure to 2.0 MPa-g resulted in a substantial increase in reforming. In contrast, at 1.0 MPa-g, the reaction resulted in a net production of CH4 due to methane production during a portion of the reaction cycle.

[0071] Additional Example - Feed Preparation for Methanol Synthesis The following is a prophetic example: A common industrial use of synthesis gas feed is the formation of methanol and / or dimethyl ether. After conversion to methanol and / or dimethyl ether, various synthetic routes are available to form larger hydrocarbons.

[0072] Some methanol synthesis processes can have problems when the syngas feed contains substantial amounts of CO. For example, syngas containing 5.0 vol.% or more of CO can result in substantial coke formation on the methanol synthesis catalyst, which can reduce the catalytic activity and / or run length of the methanol synthesis process.

[0073] Due to problems with coking, it is preferred to have 5.0 vol.% CO2 or less in the syngas feed. Unfortunately, for syngas feeds containing 5.0 vol.% CO2 or more, the options for CO2 removal without CO removal are limited. Amine washes can be used, but this also results in a substantial increase in the water content of the feed.

[0074] For synthesis gas feeds having CO2 of 5.0 vol.% or more, or 10 vol.% or more, or 20 vol.% or more, a reverse water gas shift reaction in a countercurrent reactor can be used to reduce the CO2 content in the feed to 1.0 vol.% to 5.0 vol.%, or 1.0 vol.% to 3.0 vol.%, thereby reducing or minimizing the potential for coke formation during processes using catalysts that are sensitive to CO2 content. Additionally, the ability to reduce the CO2 content to 3.0 vol.% or less contrasts with conventional synthesis gas formation methods that often result in CO2 contents of 4.0 vol.% or more.

[0075] To convert a syngas feed containing a high content of CO2 to a feed having a CO2 content of 5.0 vol.% or less (or 3.0 vol.% or less), the syngas feed can be optionally supplemented with H2, CH4 and / or H2O and then passed into the countercurrent reactor. H2 can be added so that the ratio of H2 to CO2 in the input stream to the countercurrent reactor is 2.0 or more, or 3.0 or more, or 5.0 or more. CH4 and / or H2O can be added to help shift the reforming / methanation equilibrium toward the formation of CO and H2.

[0076] Additional Embodiments Embodiment 1. A method of carrying out a reaction in a reverse flow reactor, comprising: reacting a fuel mixture comprising a fuel, 0.1 vol.% or more O2 and a diluent under combustion conditions in a combustion zone in the reverse flow reactor to form a combustion gas; and heating a conversion zone comprising a catalyst composition to an average conversion zone temperature of 400°C or more; and exposing a reactant stream comprising 5.0 vol.% to 50 vol.% CO2, 20 vol.% or more H2, with a ratio of H2 to CO2 of 2.0 or more to the catalyst composition in the conversion zone under reverse water gas shift reaction conditions to form a product stream comprising H2O and CO, wherein a direction of flow of the reactants in the reaction zone is reverse to a direction of flow of the fuel mixture.

[0077] Embodiment 2. The method of embodiment 1, wherein the reactant stream comprises a ratio of H2 to CO2 of 3.0 or greater, or the reactant stream comprises 10% CO2 by volume or greater, or a combination thereof.

[0078] Embodiment 3. The method of any of the above embodiments, further comprising mixing a syngas feed containing 10% or more by volume of CO with an H2-containing stream to form a reactant stream.

[0079] Embodiment 4. The method of any of the above embodiments, wherein the product stream comprises H2, and wherein the product stream comprises a syngas ratio of 2.0 or greater.

[0080] Embodiment 5. The method of any of the above embodiments, wherein the conversion zone is heated to an average conversion zone temperature of 500° C. or greater, or the peak conversion zone temperature at the end of reaction of the fuel mixture is 800° C. or greater, or a combination thereof.

[0081] 6. Stopping the exposure of the reactant streams; and reacting the fuel mixture under combustion conditions in the combustion zone after the exposure to heat the conversion zone to an average conversion zone temperature. 4. The method of any of the above embodiments, further comprising:

[0082] Embodiment 7. The method of any of the above embodiments, wherein the reactant stream further comprises 0.5% to 15% by volume of reformable hydrocarbons, the reformable hydrocarbons optionally comprising methane.

[0083] Embodiment 8. The method of embodiment 7, wherein the reactant stream further comprises 5.0 vol.% to 50 vol.% HO, and the ratio of HO to reformable hydrocarbons is 3.0 or greater.

[0084] Embodiment 9. The method of embodiment 7 or 8, wherein the product stream comprises reformable hydrocarbons, and the volume percentage of the reformable hydrocarbons in the product stream is less than the volume percentage of the reformable hydrocarbons in the reactant stream during the exposure step.

[0085] Embodiment 10. The reverse water gas shift conditions include a pressure of 1.4 MPa-g or more, or the reverse water gas shift conditions are 10,000 hours -1 ~50,000 hours -1 or a combination thereof.

[0086] Embodiment 11. The method of any of the above embodiments, wherein the catalyst composition comprises a reforming catalyst.

[0087] Embodiment 12. The method of embodiment 11, wherein the catalyst composition comprises Ni, Co, Fe, Pd, Rh, Ru, Pt, Ir, Cu, Ag, Au, Zr, Cr, Ti, V, W, Mo, Nb, Mn, Sr, La, or a combination thereof.

[0088] Embodiment 13. The method of any of the above embodiments, wherein the fuel comprises H2.

[0089] Embodiment 14. The method of any of the above embodiments, wherein the reactor comprises a reactor system including a recuperator section and a section including a conversion zone.

[0090] Embodiment 15. The method of any of the above embodiments, wherein the reverse water gas shift conditions include at least one of an average conversion zone temperature, and a peak conversion zone temperature at the end of reaction of the fuel mixture.

[0091] Additional embodiment A. The reverse water gas shift conditions are a pressure of 0 MPa-g to 10.3 MPa-g or 1.4 MPa-g to 10.3 MPa-g, and a reaction time of 1,000 hours. -1 ~50,000 hours -1 Or 10,000 hours -1 ~50,000 hours -1 The method of any of the above embodiments, comprising a gas hourly space velocity of

[0092] While the present invention has been described and illustrated by reference to specific embodiments, those skilled in the art will recognize that the present invention is directed to variations not necessarily illustrated herein. For this reason, reference should be made solely to the appended claims for purposes of determining the true scope of the present invention.

Claims

1. A method for carrying out a reaction in a counter-flow reactor, comprising: in a combustion zone in the counter-flow reactor, under combustion conditions, reacting a fuel with at least 0.1% by volume of O ２ and reacting a fuel mixture containing a diluent to form combustion gases and heating a conversion zone containing the catalyst composition to an average conversion zone temperature of 400° C. or greater; and reacting a fuel mixture containing 5.0 vol. % to 50 vol. % CO ２ , 20% by volume or more H ２ Including CO ２ H for ２ a ratio of H to H+ of 2.0 or greater is exposed to the catalyst composition in the conversion zone under reverse water gas shift reaction conditions to produce H ２ forming a product stream comprising O and CO, the flow direction of the reactants in a reaction zone being opposite to the flow direction of the fuel mixture; the reactant stream further comprises 0.5% to 15% by volume of a reformable hydrocarbon; The method, wherein the reverse water gas shift reaction conditions comprise a pressure of 1.4 MPa-g or greater.

2. The reactant flow is 3.0 CO ２ H for ２ or the reactant stream comprises 10 vol. % or more CO ２ 10. The method of claim 1, comprising, or a combination thereof.

3. 10% by volume or more of CO ２ A synthesis gas feed containing H ２ 3. The method of claim 1 or 2, further comprising mixing with a containing stream to form the reactant stream.

4. The product stream is H ２ and the product stream comprises a syngas ratio of 2.0 or greater.

5. 5. The method of any one of claims 1 to 4, wherein the conversion zone is heated to an average conversion zone temperature of 500°C or greater, or the peak conversion zone temperature at the end of reaction of the fuel mixture is 800°C or greater, or a combination thereof.

6. Stopping the exposure of the reactant flow; and reacting the fuel mixture under combustion conditions in the combustion zone after said exposing to heat the conversion zone to said average conversion zone temperature. The method of any one of claims 1 to 5, further comprising:

7. The method of any one of claims 1 to 6, wherein the reformable hydrocarbons optionally comprise methane.

8. The reactant stream is 5.0% to 50% by volume H ２ H for the reformable hydrocarbon further comprising O ２ 8. The method of claim 7, wherein the ratio of O is 3.0 or greater.

9. 9. The method of claim 7 or 8, wherein the product stream comprises the reformable hydrocarbons, and the volume percentage of the reformable hydrocarbons in the product stream is less than the volume percentage of the reformable hydrocarbons in the reactant stream during the exposure step.

10. The reverse water gas shift reaction conditions are 10,000 hours －１ ~50,000 hours －１ The process according to any one of claims 1 to 9, comprising a gas hourly space velocity of

11. The method of any one of claims 1 to 10, wherein the catalyst composition comprises a reforming catalyst.

12. 12. The method of claim 11, wherein the catalyst composition comprises Ni, Co, Fe, Pd, Rh, Ru, Pt, Ir, Cu, Ag, Au, Zr, Cr, Ti, V, W, Mo, Nb, Mn, Sr, La, or combinations thereof.

13. The fuel is H ２ The method according to any one of claims 1 to 12, comprising:

14. The method of any one of claims 1 to 13, wherein the reactor comprises a reactor system including a recuperator section and a section including the conversion zone.

15. The method of any one of claims 1 to 14, wherein the reverse water gas shift reaction conditions include at least one of an average conversion zone temperature and a peak conversion zone temperature at the end of reaction of the fuel mixture.

16. The reverse water gas shift reaction conditions are a pressure of 0 MPa-g to 10.3 MPa-g or 1.4 MPa-g to 10.3 MPa-g and a reaction time of 1,000 hours. －１ ~50,000 hours －１ or 10,000 hours －１ ~50,000 hours －１ The process according to any one of claims 1 to 15, comprising a gas hourly space velocity of

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