Process intensification for a countercurrent reactor

Abstract

Systems and methods are provided for improving the thermal management and / or efficiency of a reaction system including a counter-flow reactor for carrying out at least one endothermic reaction and at least one supplemental exothermic reaction. The supplemental exothermic reaction can be carried out in a recuperative zone of the counter-flow reactor system. By integrating the supplemental exothermic reaction in the recuperative zone, heat generated from the supplemental exothermic reaction can be absorbed by heat transfer surfaces in the recuperative zone. The absorbed heat can then be used to heat at least a portion of the fuel and oxidant for a combustion reaction carried out during regeneration, thereby reducing the amount of combustion required to achieve a desired temperature profile at the end of the regeneration step.

Description

Technical Field

[0001] The present invention relates to improved operation and thermal management in a countercurrent reactor.

Background Art

[0002] A countercurrent reactor is an example of a type of reactor that is useful for use in processes with cyclic reaction conditions. For example, due to the endothermic nature of the reforming reaction, it is necessary to introduce additional heat into the reforming reaction environment on a consistent basis. A countercurrent reactor can provide an efficient manner for introducing heat into the reaction environment. After a portion of the reaction cycle used for reforming or another endothermic reaction, a second portion of the reaction cycle for combustion or another exothermic reaction can be used to add heat to the reaction environment in preparation for the next reforming step. Patent Document 1 and Patent Document 2 provide examples of using a countercurrent reactor to carry out various endothermic processes in a cyclic reaction environment. Due to the high temperatures used in the endothermic process, further improvements to the countercurrent reaction system that provide increased thermal efficiency are desirable.

[0003] Patent Document 3 describes the production of synthesis gas in a countercurrent reactor by steam reforming followed by incomplete combustion of residual hydrocarbons at high temperature and pressure. In addition to providing additional synthesis gas, incomplete combustion provides heat to the reactor. In the method described in Patent Document 3, flow reversal is achieved by alternating the ends of the reactor used for the input of the reactant flows for carrying out steam reforming and incomplete combustion. The resulting synthesis gas can then be used for the production of methanol.

[0004] Patent Document 4 describes in-situ vaporizers and reheaters suitable for use in alternating flow systems such as pressure swing reformers.

Prior Art Documents

Patent Documents

[0005]

Patent Document 1

[0006] In various aspects, a method of operating a reverse flow reactor is provided. The method includes exposing at least a portion of a fuel mixture comprising fuel and at least 0.1 volume % O2 in a reheating zone of the reverse flow reactor to at least one heated surface to heat at least a portion of the fuel mixture. The method further includes reacting the fuel mixture under combustion conditions in the reheating zone to form combustion gases and heating one or more regenerative surfaces to a regenerative surface temperature in a reaction zone of the reverse flow reactor. The reaction zone can include a catalyst composition for an endothermic reaction. The reheating zone can include a catalyst composition for a supplementary exothermic reaction. The method further includes exposing a reactant stream in the reaction zone to one or more regenerative surfaces to increase the temperature of the reactant stream. The method further includes exposing the reactant stream to a catalyst composition for an endothermic reaction under endothermic reaction conditions in the reaction zone to form a product stream. The flow direction of the reactant stream in the reaction zone can be opposite to the flow direction of the fuel mixture stream. Additionally, the method includes exposing the product stream to a catalyst composition for a supplementary exothermic reaction under supplementary exothermic reaction conditions in the reheating zone to form a reacted product stream and heating at least one heated surface in the reheating zone.

[0007] In some embodiments, an example of a catalyst composition for an endothermic reaction can be a reforming catalyst. In such embodiments, the reactant stream can include a hydrocarbon that can be reformed and steam, and the product stream can include hydrogen.

[0008] Examples of catalyst compositions for complementary exothermic reactions include, but are not limited to, a water gas shift catalyst, a methanol synthesis catalyst, a methanol conversion catalyst, an olefin oligomerization catalyst, a Fischer-Tropsch catalyst for olefin production, or a combination thereof.

Brief Description of the Drawings

[0009]

Figure 1

Figure 2

Figure 3

Mode for Carrying Out the Invention

[0010] All numerical values within the detailed description and claims herein are modified by the term “about” or “approximately” and experimental errors and variations expected by those skilled in the art are considered.

[0011] In various aspects, systems and methods are provided for improving the thermal management and / or efficiency of a reaction system that includes a countercurrent reactor for performing at least one endothermic reaction and at least one supplemental exothermic reaction. The supplemental exothermic reaction is executable in a reheating zone of the countercurrent reactor system. By integrating the supplemental exothermic reaction into the reheating zone, the heat generated from the supplemental exothermic reaction can be absorbed by a heat transfer surface in the reheating zone. The absorbed heat can then be used to heat at least a portion of the fuel and oxidant for the combustion reaction performed during regeneration, thus reducing the amount of combustion required to achieve the desired temperature profile at the end of the regeneration step. Additionally, by integrating the supplemental exothermic reaction into the countercurrent reaction system, the amount of heat lost and / or the amount of heat that needs to be recovered from another heat recovery step, such as by using a heat exchanger, can be reduced or minimized.

[0012] When performing an endothermic reaction at high temperatures, a countercurrent reactor and / or other reactors having flows in opposite directions at different stages of the reaction cycle may be useful. Examples of high temperatures correspond to 600 °C or higher, or 800 °C or higher, for example up to 1600 °C or, if possible, higher temperatures. Various challenges in operating a countercurrent reactor are related to thermal management during the reaction cycle. For example, a substantial amount of energy is required to heat the input stream. Additionally, the effluent stream from the countercurrent reactor contains substantial heat and requires additional heat recovery. This includes effluent streams corresponding to products from endothermic reactions such as synthesis gas products generated during reforming or olefin and aromatic products generated during dehydrogenation reactions. Such heat recovery steps often include heat exchangers, thereby limiting the thermal efficiency.

[0013] During operation of the countercurrent reactor, the flow from the first direction corresponds to the combustion or regeneration flow and can be used to heat the reaction zone within the reactor to the desired temperature. Then, the flow in the opposite direction can be used to pass reagents for the desired endothermic reaction through the reactor. The heat stored within the reactor during the regeneration step is used to provide heat for the desired endothermic reaction. It is possible to pass the reaction products from the endothermic reaction through a reheating zone, where a portion of the heat from the reaction products is transferred to the heat transfer surface of the reheating zone. The heat stored in the reheating zone can then be used to heat at least a portion of the combustion or regeneration flow.

[0014] In various embodiments, additional heat can be recovered in the reheating zone by performing a supplementary exothermic reaction in the product discharge stream. A catalyst for the supplementary exothermic reaction can be integrated into the reheating zone of the countercurrent reaction system. By performing the supplementary exothermic reaction in the reheating zone, the heat generated by the supplementary exothermic reaction can be adsorbed by the heat transfer surface already present in the reheater. The additional adsorbed heat can then be transferred to the fuel and / or oxidant used for the combustion reaction during regeneration.

[0015] By concentrating supplementary exothermic reactions in the reheating zone, several potential advantages can be provided. First, by performing supplementary exothermic reactions in the reheating zone, the number of separate process elements that need to be maintained at high temperatures can be reduced. Each separate high-temperature process element is susceptible to energy (heat) loss to the environment, so overall thermal efficiency can be improved by reducing the number of high-temperature process elements. Another advantage is that the number of heat exchangers required to improve process efficiency can be reduced or minimized. If the supplementary exothermic reaction is carried out in a separate vessel, separate heat exchangers are required for heat recovery after both the first reactor vessel and the vessel for the supplementary exothermic reaction. In contrast, by performing the supplementary exothermic reaction in the reheating zone, a single heat exchange stage can be used for heat recovery from all of the reactions occurring in the first reactor. As a result, the supplementary exothermic reaction can be carried out before passing the reactor effluent to an external heat exchange stage. Additionally, the heat transfer occurring within the reheating zone corresponds to direct heat transfer where the material providing the heat and the material receiving the heat are in direct contact. This is in contrast to indirect heat transfer that occurs in a heat exchanger, where heat is transferred through the walls of conduits or vessels to pass heat from a first fluid to a second fluid.

[0016] The type of supplementary exothermic reaction to be carried out can be dependent on the nature of the endothermic reaction. When hydrocarbon reforming is the endothermic reaction carried out during the reaction cycle, one suitable supplementary exothermic reaction can be the high-temperature water-gas shift reaction for adjusting the ratio of H2 to CO in the syngas produced by reforming. In this discussion, hydrocarbon reforming and the water-gas shift reaction are used as an example of an endothermic reaction and a supplementary exothermic reaction to illustrate the principle of the present invention. It is understood that other convenient combinations of endothermic reactions and supplementary exothermic reactions can also be used. For example, another option can be hydrocarbon reforming as the endothermic step, combined with Fischer-Tropsch synthesis of olefins from syngas as the exothermic step. In such an embodiment, the supplementary reaction can be carried out in the presence of a suitable catalyst for the synthesis of olefins by the Fischer-Tropsch process. Suitable catalysts for the production of olefins by the Fischer-Tropsch process include, for example, iron-based Fischer-Tropsch catalysts, such as catalysts corresponding to iron particles supported on a non-reactive carrier and / or manganese-promoted iron-based catalysts. As yet another example, when alkane dehydrogenation is carried out as the endothermic step, olefin oligomerization can be carried out as the supplementary exothermic step.

[0017] In this discussion, the reaction cycle in a countercurrent reactor can include a reaction step in which an endothermic reaction is carried out, and a regeneration step in which a "first" exothermic reaction is carried out. The "first" exothermic reaction corresponds to a combustion reaction used to heat the reaction zone during the regeneration step. The supplementary exothermic reaction corresponds to an exothermic reaction carried out in a reheating zone during the reaction step.

[0018] In this discussion, unless otherwise specified, all volume ratios correspond to volume ratios shown based on the volume of the ratio quantity at standard temperature and pressure (20 °C, 100 kPa). This enables consistent specification of volume ratios even if the two combustion gas volumes being compared may exist at different temperatures and pressures. When a volume ratio is indicated for the combustion gas supplied into the reactor, it is possible to use the corresponding flow rate of the gas per unit time under standard conditions for comparison.

[0019] Process Example - Reverse Reforming and Regeneration by Integrated Aqueous Gas Shift Examples of endothermic reactions that can be carried out in a reverse reactor system are reforming of hydrocarbons under steam reforming conditions in the presence of H2O, under dry reforming conditions in the presence of CO2, or under conditions where both H2O and CO2 are present in the reaction environment. As a general overview of the operation during reforming in a swing reactor such as a reverse reactor, heat for the reactor can be provided using the regeneration step or portion of the reaction cycle. Then, reforming can occur during the reforming step or portion of the cycle, and the heat provided during the reactor regeneration step is consumed by the reforming reaction. During reactor regeneration, fuel and an oxidant are introduced into the reactor from the regeneration end of the reactor. The fuel and oxidant move from the regeneration end of the reactor towards the reaction zone and pass through the reheating zone. The layers and / or monoliths of the reheating zone of the reactor can absorb heat but typically do not contain a catalyst for reforming. However, a catalyst for an additional exothermic reaction can be included in the reheating zone. As the fuel and oxidant pass through the reheating zone, heat is transferred from the reheating zone to the fuel and oxidant. Combustion does not occur immediately, but instead, the location of combustion is controlled to occur in the central portion of the reactor. During the regeneration step, the flow of reactants continues, leading to additional transfer of heat generated from combustion into the reforming end (and thus the reaction zone) of the reactor.

[0020] After a sufficient time, the combustion reaction is stopped. Any residual combustion products and / or reactants can optionally be removed. The reforming step or portion of the reaction cycle can then be initiated. Reactants for reforming can be introduced into the reforming end of the reactor and can thus flow effectively in a direction opposite to the flow during regeneration. Layers and / or monoliths in the reforming portion of the reactor can contain a catalyst for reforming. In various embodiments, at least a portion of the catalyst can correspond to a catalyst formed from the ceramic compositions described herein. When reforming occurs, the heat introduced into the reforming region during combustion can be consumed by the endothermic reforming reaction. After exiting the reforming region, the reformed products (and unreacted reactants) are no longer exposed to the reforming catalyst. After the reformed products pass through the reheating region, heat can be transferred from the products to the regeneration region. Additionally, in embodiments where there is a catalyst for an additional exothermic reaction, the additional exothermic reaction can be carried out to generate additional heat. After a sufficient time, the reforming process can be stopped and the remaining reformed products can optionally be recovered or removed from the reactor and the cycle can be started again by the regeneration step.

[0021] The reforming reaction carried out in the reactor can correspond to reforming methane and / or other hydrocarbons using steam reforming in the presence of H2O, dry reforming in the presence of CO2, or "bi" reforming using both H2O and CO2. Examples of the stoichiometry for the steam, dry, and "bi" reforming of methane are shown in equations (1)-(3). (1) Dry reforming: CH4 + CO2 = 2CO + 2H2 (2) Steam reforming: CH4 + H2O = CO + 3H2 (3) Bi reforming: 3CH4 + 2H2O + CO2 = 4CO + 8H2.

[0022] As shown in Equations (1)-(3), dry reforming can produce a lower ratio of H2 to CO than steam reforming. A reforming reaction carried out with only steam can generally produce a ratio of H2 to CO of about 3, such as 2.5 - 3.5. In contrast, a reforming reaction carried out in the presence of CO2 can produce a lower ratio of H2 to CO, probably a ratio of about 1.0 or less. By using a combination of CO2 and H2O during reforming, the reforming reaction can potentially be controlled to produce a wide range of ratios of H2 to CO in the resulting synthesis gas.

[0023] The reforming reactions shown in Equations (1)-(3) are endothermic reactions. One of the challenges in reforming on a commercial scale may be to provide heat in an efficient manner while reducing or minimizing the introduction of additional components into the desired synthesis gas product. A periodic reaction system, such as a countercurrent reactor system, can provide heat in a desirable manner by having a cycle that includes a reforming step and a regeneration step. During the regeneration step, combustion can be carried out within a selected region of the reactor. The gas flow during regeneration can assist in transferring this heat from the combustion region to an additional portion of the reforming region within the reactor. The reforming step within the cycle can be a separate step so that the products from combustion can be reduced or minimized from being incorporated into the reactants and / or the products from reforming. The reforming step can consume heat and thereby reduce the temperature of the reforming region.

[0024] It should be noted that the ratio of H2 to CO in the synthesis gas may also depend on the water-gas shift equilibrium. The stoichiometry of Equations (1)-(3) shows ratios of about 1 or about 3 for dry reforming and steam reforming, respectively, but the equilibrium amounts of H2 and CO in the synthesis gas may differ from the reaction stoichiometry. The equilibrium amounts can be determined based on the water-gas shift equilibrium, which is related to the concentrations of H2, CO, CO2, and H2O based on Reaction (4) H2O + CO <=> H2 + CO2.

[0025] Most reforming catalysts such as rhodium and / or nickel can also be useful as water-gas shift catalysts. Therefore, when H2O and / or CO2 are also included in the reaction environment for producing H2 and CO, the initial stoichiometry from the reforming reaction may be changed based on the water-gas shift equilibrium. However, this equilibrium is also temperature-dependent, and higher temperatures are preferred for the production of CO and H2O. As a result, the ratio of H2 to CO that occurs when forming syngas is suppressed by the water-gas shift equilibrium at the temperature in the reaction zone when the syngas is produced.

[0026] To further change the ratio of CO to H2O in the syngas, another high-temperature water-gas shift catalyst can be included in the reheating zone of the countercurrent reactor. An example of a high-temperature water-gas shift catalyst is a catalyst containing a mixture of iron oxide and chromium oxide. As described above, the reforming catalysts described herein can also provide water-gas shift activity in a high-temperature environment. The reheating zone can be at a lower temperature than the reaction zone so that an increase in the amount of H2 relative to CO is obtained by the water-gas shift reaction. For example, the temperature of the reaction zone can be 600 °C or higher, while the temperature of the reheating zone can be 300 °C to 450 °C. Optionally, the equilibrium can be further shifted towards the production of H2 by increasing the amount of water present in the diluent and / or decreasing the amount of CO2 in the diluent.

[0027] Several advantages can be achieved by performing a supplementary water gas shift reaction (or another supplementary exothermic reaction) in the reheating zone. First, as described above, by performing the water gas shift at a lower temperature than the reaction zone, a potentially more favorable ratio of H2 to CO can be achieved in the syngas. Second, additional heat can be recovered for use in subsequent regeneration steps. In the water gas shift reaction, the formation of H2 and CO2 from H2O and CO is exothermic. Thus, the reheating zone can adsorb heat released by the exothermic water gas shift reaction in addition to adsorbing heat from the syngas product. By increasing the heat available in the reheating zone, a corresponding reduction in the amount of fuel burned during regeneration becomes possible. Third, performing the water gas shift reaction in the reheating zone means that additional vessel and / or reactor volume is not maintained at high temperatures. This results in a reduced footprint for the reaction system and also reduced or minimized heat loss to the environment.

[0028] One common source for methane is natural gas. In some applications, natural gas, including associated hydrocarbons and impurity gases, may be used as a feed for reforming reactions. The natural gas supplied may be sweetened and / or dehydrated natural gas. Natural gas generally contains associated gases such as ethane and other alkanes in various concentrations, preferably at concentrations lower than methane. The natural gas supplied may contain impurities such as H2S and nitrogen. More generally, the hydrocarbon feed for reforming can include any convenient combination of methane and / or other hydrocarbons. Optionally, the reforming feed may contain some hydrocarbon-based compounds such as alcohols or mercaptans. Hydrocarbon-based compounds are similar to hydrocarbons but contain one or more heteroatoms different from carbon and hydrogen. In some embodiments, additional components present in the feed can correspond to impurities such as sulfur that can be adsorbed onto the catalyst monolith during a reduction cycle such as a reforming cycle. Such impurities can then be oxidized in a subsequent cycle to form sulfur oxides, which can then be reduced to release additional sulfur-containing components (or other impurity-containing components) into the reaction environment.

[0029] In some embodiments, the feed for reforming contains 5 wt% or more, or 10 wt% or more, or 15 wt% or more, or 20 wt% or more, for example up to 50 wt% or, if possible, even higher wt% of C such as ethane or propane, compared to the total weight of the hydrocarbons in the feed for reforming. 2+It can contain compounds. Note that nitrogen and / or other gases that are non-reactive in a combustion environment such as H2O and CO2 may also be present in the feed for reforming. In embodiments where the reformer corresponds to an on-board reforming environment, such non-reactive products can be optionally introduced into the feed, for example, based on the recycling of exhaust gas into the reformer. Moreover, or alternatively, the feed for reforming can contain 40 wt% or more, or 60 wt% or more, or 80 wt% or more or 95 wt% or more of methane, and for example, the feed can be substantially composed of methane (98 wt% or more). In embodiments where the reforming corresponds to steam reforming, the molar ratio of steam molecules to carbon atoms in the feed can be 0.3 to 4.0. Note that ethane has two carbon atoms per molecule, while methane has one carbon atom per molecule. In embodiments where the reforming corresponds to dry reforming, the molar ratio of CO2 molecules to carbon atoms in the feed can be 0.05 to 3.0.

[0030] Within the reforming zone of the countercurrent reactor, the temperature can vary across the zone due to the nature of how heat is applied to the reactor and / or the kinetics of the reforming reaction. The highest temperature portion of the zone can typically be found near the central portion of the reactor. This central portion can be described as the mixing zone where combustion begins during regeneration. At least a portion of the mixing zone can correspond to a portion of the reforming zone if a monolith having a reforming catalyst extends into the mixing zone. As a result, the location where combustion begins during regeneration can typically be close to the end of the reforming zone within the reactor. Note that the position of the combustion catalyst within the reactor can overlap with the position of the reforming catalyst within the reactor such that some portions of the reactor can correspond to both the combustion zone and the reaction zone. The temperature can decrease from the center of the reactor towards the end of the reactor. As a result, the temperature at the start of the reforming zone (at the end of the reactor) can be cooler than the temperature at the end of the reforming zone (at the central portion of the reactor).

[0031] When a reforming reaction occurs, the temperature in the reforming zone can decrease. The rate of temperature decrease can be related to the amount of hydrocarbon available for reforming and / or the kinetic factors of the temperature at a given location in the reforming zone. As the reforming feed moves to the end of the reforming zone, it is possible to consume the reactants of the feed, thereby reducing the amount of reforming occurring at downstream locations. However, an increase in the temperature of the reforming zone as the reactants move across the reforming zone can lead to an increase in the reaction rate.

[0032] At about 500 °C, it is possible to sufficiently reduce the reaction rate for reforming, and little or no additional reforming will occur. As a result, in some embodiments, as the reforming reaction proceeds, the starting portion of the reforming zone can be sufficiently cooled to effectively stop the reforming reaction within a portion of the reforming zone. This makes it possible to move the location where reforming starts in the reactor to a location further downstream relative to the start of the reforming zone. If a sufficient portion of the reforming zone has a temperature below 500 °C, or below 600 °C, it is possible to stop the reforming step during the reaction cycle and perform regeneration. Alternatively, based on the amount of heat introduced into the reactor during regeneration, the amount of heat consumed (and heat loss to the environment) during reforming can be balanced with the amount of heat added during regeneration based on the amount of reaction time, and the reforming portion of the reaction cycle can be stopped. After the reforming process is stopped, any remaining syngas product still in the reactor can optionally be recovered before starting the regeneration step of the reaction cycle.

[0033] Next, the regeneration process can be initiated. During regeneration, fuels such as methane, natural gas, or H2 and oxygen can be introduced into the reactor and combusted. The location where the fuel and the oxidant can be mixed can be controlled by any convenient method, such as introducing the fuel and the oxidant through separate channels. During regeneration, the non-reformed end of the reactor can be maintained at a lower temperature by delaying combustion until the reactants reach the central part of the reactor. This allows a temperature peak to be obtained in the central part of the reactor. The temperature peak can be within the range of a portion of the reactor that also includes the reforming catalyst. During the regeneration cycle, the temperature within the reforming reactor can be increased sufficiently to enable reforming during the reforming part of the cycle. This allows a peak temperature within the reactor of 1100 °C or higher, or 1200 °C or higher, or 1300 °C or higher, or potentially even higher temperatures to be obtained.

[0034] The relative lengths of time and the reactant flow rates for the reforming and regeneration portions of the process cycle can be selected such that a balance is achieved between the heat provided during regeneration and the heat consumed during reforming. For example, one option can be to select a reforming step having a length similar to that of the regeneration step. Based on the flow rates of hydrocarbons, H2O, and / or CO2 during the reforming step, the endothermic heat requirement for the reforming reaction can be determined. This heat requirement 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 reforming 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 such that a balance is achieved between the heat provided by regeneration and the heat consumed during reforming.

[0035] In addition to providing heat, the reactor regeneration step during the reaction cycle can also enable coke removal from the catalyst in the reforming 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 hydrocarbon introduced into the reforming zone can form coke instead of forming CO or CO2. This coke can potentially interfere with access to the catalyst sites (such as metal sites) of the catalyst. In some embodiments, the formation rate can increase in portions of the reforming zone that are exposed to higher temperatures, such as portions of the reforming zone that are 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 increases. At the temperature reached during regeneration, at least a portion of the coke that forms during reforming can be removed as CO or CO2.

[0036] Due to temperature fluctuations within the reactor, several options can be used to characterize the temperature within the reactor and / or within the reforming 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, determination of the temperature within the reactor requires the presence of a measurement device such as a thermocouple. Rather than attempting to measure the temperature within the reforming zone, the average (bed or monolith) temperature within the reforming zone can be defined based on the average of the temperature at the start of the reforming zone and the temperature at the end of the reforming zone. Another option can be to characterize the peak temperature within the reforming zone after the regeneration step of the reaction cycle. Generally, the peak temperature can occur at or near the end of the reforming zone and can depend on the location where combustion starts within 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 reforming step. Such a temperature difference can be characterized at the location of the peak temperature within the reactor, the inlet to the reforming zone, the outlet from the reforming zone, or any other convenient location.

[0037] In various embodiments, the reaction conditions for hydrocarbon reforming include an average reforming zone temperature in the range of 400 °C to 1200 °C (or higher); a peak temperature within the reforming zone of 800 °C to 1500 °C; a temperature difference at the peak temperature location between the end of the regeneration step and the end of the subsequent reforming step of 25 °C or more, or 50 °C or more, or 100 °C or more, or 200 °C or more, for example up to 800 °C, or even higher if possible; a temperature difference at the inlet to the reforming zone between the end of the regeneration step and the end of the subsequent reforming step of 25 °C or more, or 50 °C or more, or 100 °C or more, or 200 °C or more, for example up to 800 °C, or even higher if possible; and / or a temperature difference at the outlet from the reforming zone between the end of the regeneration step and the end of the subsequent reforming step of 25 °C or more, or 50 °C or more, or 100 °C or more, or 200 °C or more, for example up to 800 °C, or even higher if possible. One or more of these can be included.

[0038] Regarding the average reforming zone temperature, in various embodiments, the average temperature for the reforming zone can be 500 °C to 1500 °C, or 400 °C to 1200 °C, or 800 °C to 1200 °C, or 400 °C to 900 °C, or 600 °C to 1100 °C, or 500 °C to 1000 °C. Additionally, or alternatively, regarding the peak temperature of the reforming zone (which perhaps corresponds to a location in the reforming zone close to the location for combustion of the regeneration reactants), the peak temperature can be 800 °C to 1500 °C, or 1000 °C to 1400 °C, or 1200 °C to 1500 °C, or 1200 °C to 1400 °C.

[0039] Additionally, or alternatively, the reaction conditions for hydrocarbon reforming include a pressure of 0 psig to 1500 psig (10.3 MPa), or 0 psig to 1000 psig (6.9 MPa), or 0 psig to 550 psig (3.8 MPa); and -1 ~50,000 hours -1 a gas hourly space velocity of the reforming reactants. The space velocity corresponds to the volume of the reactants per unit volume of the monolith per unit time. The volume of the monolith is defined as the volume of the monolith when it is a solid cylinder.

[0040] In some embodiments, the advantage of operating the reforming reaction at elevated temperatures can be the ability to convert substantially all of the methane and / or other hydrocarbons in the reforming feed. For example, for reforming processes in which water is present in the reforming environment (i.e., steam reforming or bi-reforming), the reaction conditions can be appropriate for the conversion of 10 wt% to 100 wt%, or 20 wt% to 80 wt%, or 50 wt% to 100 wt%, or 80 wt% to 100 wt%, or 10 wt% to 98 wt%, or 50 wt% to 98 wt% of the methane in the reforming feed. Additionally or alternatively, the reaction conditions can be appropriate for the conversion of 10 wt% to 100 wt%, or 20 wt% to 80 wt%, or 50 wt% to 100 wt%, or 80 wt% to 100 wt%, or 10 wt% to 98 wt%, or 50 wt% to 98 wt% of the hydrocarbons in the reforming feed.

[0041] In other embodiments, for reforming processes in which carbon dioxide is present in the reforming environment (i.e., dry reforming or bi-reforming), the reaction conditions can be appropriate for the conversion of 10 wt% to 100 wt%, or 20 wt% to 80 wt%, or 50 wt% to 100 wt%, or 80 wt% to 100 wt%, or 10 wt% to 98 wt%, or 50 wt% to 98 wt% of the methane in the reforming feed. Additionally or alternatively, the reaction conditions can be appropriate for the conversion of 10 wt% to 100 wt%, or 20 wt% to 80 wt%, or 50 wt% to 100 wt%, or 80 wt% to 100 wt%, or 10 wt% to 98 wt%, or 50 wt% to 98 wt% of the hydrocarbons in the reforming feed.

[0042] In some other embodiments, the reforming reaction can be carried out using CO2 as a reagent under dry reforming conditions where the amount of H2O in the reaction environment is reduced or minimized. In such other embodiments, it is possible for the goal of the reforming reaction to be to produce syngas having a ratio of H2 to CO of 1.0 or less. In some embodiments, the temperature during reforming can correspond to the temperature ranges described for steam reforming. Optionally, in some embodiments, the dry reforming reaction can be carried out at a low temperature between 500°C and 700°C, or between 500°C and 600°C. In such embodiments, the ratio of H2 to CO can be 0.3 to 1.0, or 0.3 to 0.7 or 0.5 to 1.0. Carrying out the dry reforming reaction under these conditions can lead to substantial coke formation, which may require removal during regeneration to maintain catalyst activity.

[0043] Figure 1 shows an example of an arrangement for using a countercurrent reactor as part of a reaction system for hydrogen production. In the example shown in Figure 1, for clarity, a single countercurrent reactor and a single hydrogen recovery stage are shown. It is understood that any convenient number of countercurrent reactors can be used with any convenient number of hydrogen recovery stages so that continuous production of hydrogen can be carried out, even if the countercurrent reactor and the hydrogen recovery stage can have separate reaction cycles including a regeneration step.

[0044] In the example of the arrangement shown in FIG. 1, a countercurrent reactor 110 is used to carry out steam reforming as an endothermic reaction. The feed 101 for the endothermic reaction corresponds to methane and steam. The feed 101 is introduced into the countercurrent reactor 110 at the end of the reactor closest to the reaction zone 120. During the reaction (reforming) step of the reaction cycle, the feed is reformed to produce hydrogen and carbon monoxide. The reformed effluent passes through a reheating zone 130 before exiting the reactor. In the reheating zone 130, the reformed effluent is exposed to a high-temperature water gas shift catalyst that increases the relative ratio of hydrogen to carbon monoxide in the effluent. The shifted effluent 135 is then discharged from the reactor 110. The shifted effluent 135 then passes through a heat exchanger stage 140 where heat is recovered. For example, heat can be recovered as steam 148. The cooled effluent 145 can then be passed through a hydrogen recovery stage 150, such as a pressure swing adsorber, to enable the generation of a high-purity hydrogen stream 155.

[0045] After a certain period of time, the production of the hydrogen stream 155 can be stopped and the countercurrent reactor 110 and the hydrogen recovery stage 150 can be regenerated. In the ideal example shown in FIG. 1, the regeneration step for the hydrogen recovery stage 150 can be carried out simultaneously with the regeneration step for the countercurrent reactor 110. In other embodiments, a valve configuration can be used to decouple the reaction cycle for one or more hydrogen recovery stages 150 from the reaction cycle for one or more countercurrent reactors.

[0046] During the regeneration step, a purge stream 152 is introduced into a hydrogen recovery stage (e.g., a pressure swing adsorber) 150 to produce a tail gas 142. The tail gas can contain, for example, CO and unreacted methane present in the reformate effluent. Additional fuel 107 can be added to the tail gas 142. Optionally, a recycle stream 133 of CO2 and / or H2O can also be added as a diluent gas for the regeneration step. Then, the combined stream of the tail gas 142, the additional fuel 107, and the optional recycle stream 133 can be passed through the reactor 110 near the end closer to the reheating zone 130. This combined stream is heated in the reheating zone 130 and combusted at or near the interface between the reheating zone 130 and the reaction zone 120. A portion of the resulting combustion effluent 122 discharged from the reactor 110 can be used to form the optional recycle stream 133.

Example

[0047] Example of a countercurrent reactor configuration For endothermic reactions carried out at high temperatures such as hydrocarbon reforming, a countercurrent reactor can provide a suitable reaction environment to supply heat to the endothermic reaction.

[0048] In a countercurrent reactor, the heat required for the endothermic reaction may be provided by creating a high-temperature heat bubble in the center of the reactor. Then, a two-step process can be used where the heat is (a) added to the reactor bed or monolith via in-situ combustion and then (b) removed via an endothermic process such as reforming, pyrolysis, or steam cracking. This type of configuration can provide the ability to consistently manage and confine the high-temperature bubble in a reactor region that can withstand such conditions for an extended period. The countercurrent reactor system enables the first endothermic and regeneration processes to be carried out in a substantially continuous manner.

[0049] In some embodiments, the countercurrent reactor system can correspond to a single reactor having a reaction zone and a reheating zone. In other embodiments, the countercurrent reactor system can include first and second reactors arranged in series with respect to a common flow path and, optionally but preferably, along a common axis. The common axis can be horizontal or vertical.

[0050] During the regeneration step, reactants (e.g., fuel and oxygen) can be combined or mixed in the reaction zone in order to combust them in situ therein and to create a high temperature region or a thermal bubble within the central portion of the reaction system. The thermal bubble can correspond to a temperature that is at least approximately the initial temperature of the endothermic reaction. Typically, as heat is transferred from the thermal bubble in the central portion of the reactor towards the end of the reactor, the temperature decreases, so the temperature of the thermal bubble can be higher than the initial temperature for the endothermic reaction. In some embodiments, the combining can be enhanced by a reactant mixer that mixes the reactants by means of a mixer that is optionally positioned between the first and second reactors to promote substantially complete combustion / reaction at a desired location. The combustion process is executable for a long enough period to replace, in a preferably not all but a substantial portion of the heat (e.g., thermal bubble) generated by the reaction into the second reactor and at least partially through it, and to serve to reduce or minimize waste heat and heating of the second reactor. This heat is transferred, for example, to one or more surfaces of the second reactor and / or of the reaction zone for the endothermic reaction of the reactor. The combustion gas may be discharged through the second reactor, but preferably most of the heat is retained within the second reactor. The amount of heat replaced into the second reactor during the regeneration step can also be limited or determined by the desired exposure time or space velocity that the hydrocarbon feed gas has in the endothermic reaction environment. In embodiments where a single reactor is used, the heat generated by the reaction can be replaced within and / or at least partially through the reaction zone of the reactor, but preferably the replacement can also reduce or minimize waste heat for the discharge of the heated gas from the reactor.

[0051] After regenerating or heating the second reaction medium (which can include one or more surfaces and / or be equivalent thereto), in the next / reverse step or cycle, can reactants for the endothermic reaction, methane (and / or natural gas and / or another hydrocarbon), be fed into the second reactor or can it be flowed in from a direction opposite to the flow direction during the heating step? For example, in a reforming process, can methane (and / or natural gas and / or another hydrocarbon) be fed into the second reactor or can it be flowed in? Methane can be contacted with the heated second reactor and the mixing machine medium in the thermal bubble region to transfer heat for the reaction energy to the methane.

[0052] For some embodiments, the basic two-step asymmetric cycle of the countercurrent regenerative bed reactor system relates to a reactor system having two regions / reactors; a first or reheater / quench region (7) and a second or reaction region (1), shown in FIGS. 2A and 2B of FIG. 2. Both the reaction region (1) and the reheating region (7) can contain a regenerative monolith formed from a doped ceramic composition and / or other regenerative structures. The regenerative monolith or other regenerative structures, as used herein, include materials that are effective in storing and transferring heat and effective in carrying out chemical reactions. The regenerative monolith and / or other structures can correspond to any convenient type of material that is suitable for storing heat, transferring heat, and catalyzing the reaction. Examples of structures include layered or packed material ceramic beads or spheres, ceramic honeycomb materials, ceramic tubes, extruded monoliths, etc., provided that they are capable of withstanding long-term exposure to temperatures higher than 1200°C, or higher than 1400°C, or higher than 1600°C such that they maintain integrity, functionality, and allow for some operating margin. In some embodiments, a catalytic ceramic monolith and / or other catalytic ceramic structures can be used without the presence of an additional washcoat.

[0053] To facilitate the description of FIG. 2, a reactor for the reforming reaction is described herein. It is understood that other convenient types of endothermic reactions can generally be carried out using a countercurrent reactor such as the reactor shown in FIG. 2.

[0054] As shown in FIG. 2B of FIG. 2, at the start of the "reaction" step of the cycle, the second end 5 of the reaction zone 1 (also described herein as the second reactor) can be at a higher temperature compared to the first end 3 of the reaction zone 1, and at least a portion of the reheater or quench zone 7 (including the first end 9) (also described herein as the first reactor) can be at a lower temperature than the reaction zone 1 to provide a quench effect to the resulting product. In the manner in which the reactor is used to perform countercurrent reforming, a methane-containing reactant feed (or other hydrocarbon-containing reactant feed) is introduced via conduit 15 to the first end 3 of the reforming or reaction zone 1. In various embodiments, the hydrocarbon-containing reactant feed can also contain H2O, CO2, or a combination thereof.

[0055] The flow of the feed from inlet 15 can absorb heat from the reaction zone 1, undergo an endothermic reaction, and produce the desired syngas product. As this step progresses, a shift in the temperature profile 2 can be created based on the heat transfer characteristics of the system, as indicated by the arrow. If the ceramic catalyst monolith / other catalyst structure is designed with appropriate heat transfer performance, this profile can have a relatively sharp temperature gradient, and such a gradient can move across the entire reaction zone 1 as the reforming step progresses. In some embodiments, a sharper temperature gradient profile results in an improvement in the control of the reaction conditions. In the manner in which another type of endothermic reaction is carried out, a similar shift in the temperature profile can occur such that the temperature gradient moves across the entire reaction zone 1 as the reaction step progresses.

[0056] The effluent from the reforming reaction, which can contain unreacted feed components (hydrocarbons, H2O, CO2) as well as syngas components, exits the reaction zone 1 at a high temperature through the second terminal 5, passes through the reheater 7, and then through the second terminal 11 and can exit through the first terminal 9. The reheater 7 can initially be at a lower temperature than the reaction zone 1. When the product from the reforming reaction (and optionally unreacted feed) passes through the reheating zone 7, the gas can be quenched or cooled at the first terminal 9 to a temperature near that of the reheating zone, which in some embodiments is substantially the same temperature as the regenerated feed introduced through conduit 19 into the reheater 7 during the second step of the cycle. When the reforming effluent is cooled in the reheating zone 7, a temperature gradient 4 can be created in the regenerative bed of the zone and can move throughout the reheating zone 7 during this step. The reheater 7 can be heated by the quench, which is cooled again in the second step, providing for another quench execution, and can 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 reheater at 9 through conduit 17 and can be processed for separation and recovery of various components. In some embodiments, a catalyst for an exothermic reaction supplementary to the reheating zone 7 can also be included. In such embodiments, at least a portion of the additional heat generated by the supplementary exothermic reaction can also be transferred to the regenerative bed of the reheating zone 7. For example, the catalyst can correspond to a high-temperature water-gas shift catalyst such that an exothermic water-gas shift reaction can occur in the reheating zone. Examples of high-temperature water-gas shift catalysts include catalysts based on iron oxide and / or chromium oxide.

[0057] Next, the second step of the cycle, called the regeneration step, can be initiated by reintroducing the first and second regeneration reactants through conduit 19. The first and second reactants can be passed separately through the reheater 7 and heated towards the second end 11 of the reheater 7. In the reheater 7, they can be combined for an exothermic reaction or combustion in or near the central region 13 of the reactor system.

[0058] An example of the regeneration step is shown in FIG. 2B of FIG. 2. Regeneration can involve the transfer of a significant amount of heat recovered from the reheating region 7 to the reaction region 1 to thermally regenerate the reaction layer 1 for subsequent reaction cycles. The regeneration gas / reactant can enter the reheating region 7 via, for example, conduit 19, pass through the reheating region 7, and flow into the reaction region 1. At that time, as indicated by the arrows on the exemplary graph in FIG. 2B, the temperature gradients 6 and 8 can move across the layer and are opposite but similar to the graph of the temperature gradients developed during the reaction cycle in FIG. 2A of FIG. 2. The fuel and oxidant reactants can combust in the region near the interface 13 between the reheating region 7 and the reaction region 1. The heat recovered from the reheating region, together with the heat of combustion, can be transferred to the reaction region and thermally regenerate the regeneration reaction monolith and / or layer 1 disposed therein.

[0059] In some embodiments, some conduits within the channel can carry a mixture of the first and second reactants, at least in part, for 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 kept low enough so that most of the stoichiometrically reactive reactants do not react until exiting the second end of the first reactor. The axial position of initiation of a combustion or exothermic reaction carrying the reactant mixture in those conduits can be controlled by a combination of temperature, time, and fluid dynamics. Fuel and oxygen typically require a temperature - and mixture - dependent auto - ignition time to burn. However, some reactions can occur within the axial portion of the conduit carrying the reactant mixture. However, the number of channels having such reactions is small enough that this reaction is acceptable because only an acceptable or non - significant level of impact on the overall heat balance within the reactor is present. The details of a particular reactor design can be chosen to avoid reactant mixing in as many conduits as reasonably possible.

[0060] Figure 3 shows another exemplary reactor system that may be suitable for several applications to control and delay the combustion of fuel and oxidant in order to achieve efficient regenerative heat. Figure 3 shows a single reactor system operating in a regeneration cycle. The reactor system may be considered to include two reactor regions. The reheater 27 is a region where quenching mainly occurs and a substantially isolated flow path or channel is provided for both quenching reaction gases to be transmitted through the reactor medium without causing combustion until the gas arrives near or inside the reactor core of Figure 2. In some embodiments, such a channel can also reduce or minimize the exposure of fuel and / or oxidant to the catalyst for the supplementary exothermic reaction. The reformer 2 can be a reactor where regenerative heating and methane (and / or hydrocarbon) reforming mainly occur and can be considered as a second reactor for the purposes of this specification. The first and second reactors of the reactor system are specified as separately distinguishable reactors, but the first and second reactors can be manufactured, provided, or combined in other ways in a common single reactor layer, whereby the reactor system may be described as including only a single reactor integrating both cycles within the reactor. It is understood that "the first reactor" and "the second reactor" can simply mean each region within the reactor system where respective steps such as regeneration, reforming, quenching, etc. occur and do not require separate components to be utilized for the two reactors. However, various embodiments can include a reactor system whereby the reheater reactor includes the conduits and channels described herein and the reformer reactor can similarly possess conduits. Additionally, or alternatively, some embodiments can include a reformer reactor layer that is disposed separately from and may even include materials different from the reheater reactor layer.

[0061] As discussed above, the first reactor or reheater 27 can include various gas conduits 28 for separately flowing two or more gases after entering the first end 29 of the reheater 27 and through a regeneration layer disposed therein. The first gas 30 can enter the first ends of the plurality of flow conduits 28. In addition to providing flow channels, the conduits 28 can also include an effective flow barrier (e.g., one that functions effectively as a conduit wall) to prevent cross-flow and mixing between the first and second reactants and to maintain the major reactants in a state of being effectively separated from each other until mixing is tolerated. As discussed above, each of the first and second channels can include a plurality of channels or flow paths. The first reactor can also include a plurality of substantially parallel flow segments, each including a separate first and second channel.

[0062] In some embodiments, the reheater can, as described above, be composed of one or more extruded honeycomb monoliths. Each monolith can 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 can include one or more channels for each reactant, and one or more groups of channels or conduits are for flowing one or more flows of one reactant, while the remaining portions of the conduits are for flowing one or more flows of the other reactant. It is recognized that a number of conduits can carry a mixture of the first and second reactants at the interface between the channels, but this number of conduits is proportionally small.

[0063] In an embodiment where a monolith is used, the monolith can have any convenient shape suitable for use as a catalyst surface. An example of a monolith can be an extruded honeycomb monolith. A honeycomb monolith can be an extruded structure that includes a plurality of (e.g., a plurality meaning two or more) small gas flow channels or conduits arranged in parallel in a pattern having thin walls therebetween. A small reactor can include a single monolith, but a larger reactor can include a number of monoliths, and an even larger reactor can be substantially filled with an arrangement of a number of honeycomb monoliths. Each monolith can be formed by extruding a monolith block having a shaped (e.g., square or hexagonal) cross-section and stacking such blocks two-dimensionally or three-dimensionally on top of, behind, and beside each other. Monoliths can be attractive as an internal reactor structure because they provide high heat transfer performance with minimal pressure drop.

[0064] In some embodiments, the honeycomb monolith can be characterized by having an open frontal area (or geometric void volume) of 25% - 55%, a pore or cell density (CPSI) of 50 - 2000 per square inch, or 100 - 900 cells per square inch, or 100 cells per square inch to 600 cells per square inch. For example, in one embodiment, the conduits can have a diameter / characteristic cell side length of only a few millimeters of about 1 millimeter. A reactor media component such as a monolith or another layer media is based on the volume of the first reactor used to transport reactants, 50 feet -1 ~ 3000 feet -1 (about 0.16 km -1 ~ about 10 km -1 ), or 100 feet - 1 to 2500 feet -1 (about 0.32 km -1 ~ about 8.2 km -1 ), or 200 feet -1 ~ 2000 feet -1 (about 0.65 km -1 ~ about 6.5 km -1) It is possible to provide channels that include the filling of the average wet surface area per unit volume within the range. These relatively high surface areas per unit volume value can help achieve a relatively rapid change in temperature in the reactor, as generally indicated by a relatively steep gradient in the exemplary temperature gradient profile graphs shown in FIGS. 2A or 2B of FIG. 2.

[0065] The reactor media component includes a filling having a high volumetric heat transfer coefficient (e.g., 0.02 calories / cm 3 second °C or more, or 0.05 calories / cm 3 second °C or more, or 0.10 calories / cm 3 second °C or more), has a low flow resistance (low pressure drop); has an operating temperature range that coincides with the maximum temperature occurring during regeneration; has a high thermal shock resistance; and / or has a high volumetric heat capacity (e.g., 0.10 calories / cm 3 second °C or more, or 0.20 calories / cm 3 second °C or more). Similar to the high surface area values, these relatively high volumetric heat transfer coefficient values and / or other properties can help achieve a relatively rapid change in temperature in the reactor, as generally indicated by a relatively steep gradient in the exemplary temperature gradient profile graphs in FIGS. 2A and 2B of FIG. 2. The listed values are averages based on the volume of the reactor used for the conveyance of the reactants.

[0066] In various embodiments, an appropriate heat transfer rate can be characterized by a heat transfer parameter (ΔTHT) at less than 500 °C, or less than 100 °C, or less than 50 °C. The parameter ΔTHT, as used herein, is the ratio of the layer average volumetric heat transfer rate required for reheating to the volumetric heat transfer coefficient (hv) of the layer. A volumetric heat transfer rate sufficient for reheating (e.g., calories / cm 3 second) can be calculated as the product of the gas flow rate (e.g., g / second), 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 this amount is divided by the volume of the reactor (or a portion of the reactor) through which the gas moves (e.g., cm 3) can be divided by. The volumetric heat transfer coefficient (hv) of the layer is related to the area-based coefficient (e.g., calories / cm 2 second °C), and is typically calculated as the product of the surface area for heat transfer (av, e.g., cm 2 / cm 3 ), which is also often described as the wet surface area of the packing.

[0067] In some embodiments, a washcoat can be added to the formed, sintered ceramic composition before exposing the composition to a reducing environment to form dopant metal particles. The washcoat enables the sintered ceramic composition to be impregnated with additional catalytic metal. Such additional catalytic metal can be the same as or different from the dopant metal.

[0068] One option for incorporating additional catalytic metal into the washcoat can be, for example, to impregnate the catalyst support with the additional catalytic metal by impregnation by incipient wetness. The impregnation can be carried out in an aqueous solution of a suitable metal salt or other catalytic metal precursor such as tetrammineplatinum nitrate or rhodium nitrate hydrate. The impregnated support can then be dried and / or calcined for the decomposition of the catalytic metal precursor. Various temperature profiles can potentially be used for the heating step. One or more initial drying steps such as heating at a temperature of 100°C to 200°C for 0.5 hour to 24 hours can be used to dry the support. Depending on the nature of the impregnated catalytic metal compound, the calcination for decomposing the catalytic metal precursor compound can be at a temperature of 200°C to 800°C for 0.5 hour 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.

[0069] In another embodiment, the channel conduit / flow path thereby includes, but is not limited to, a more tortuous path (e.g., sinuous, complex, helical and / or twisted, but not linear or tubular) such as a labyrinthine and diverse flow path, conduit, tube, slot and / or pore structure having channels passing through a portion of the reactor, and may include other suitable means, such as along the outer surface of the segment or within the sub-segment, to substantially maintain the first and second reactant gases in a substantially separated state from each other while passing axially through the reheater 27, to prevent cross-flow between the barrier portion and / or reactant gases that are substantially impermeable to the gas, and to maintain the first and second reactant gases in a substantially separated state from each other. A non-monolithic reaction medium may be used. Such other types of reactor media may be suitable as long as at least a portion of such media can be formed by firing the ceramic catalyst composition described herein and subsequently exposing such media to reducing conditions to activate the catalyst. For such embodiments, the complex flow path can result in an extended effective flow path, increased surface area, and improved heat transfer. Such a design may be preferred for reactor embodiments having a relatively short axial length through the reactor. A longer reactor length in the axial direction may experience an increase in pressure drop across the reactor. However, for such embodiments, the porous and / or permeable medium may include, for example, at least one of a packed bed, an array of tiles, a permeable solid medium, a substantially honeycomb-type structure, a fiber array, and a mesh-type lattice structure.

[0070] In some embodiments, the reverse flow reactor can include several types of devices or methods for directing one stream of reactants to a selected portion of the conduit. In the exemplary embodiment of FIG. 3, the gas distributor 31 can direct a second gas stream to a second gas flow channel that is substantially isolated from or not in fluid communication with the first gas channel, shown herein as channel 33. As a result, at least a portion of the gas stream 33 can be kept isolated from the gas stream 30 during axial movement of the reheater 27. In some embodiments, the regenerative bed and / or monolith of the reheating zone can include channels having a gas or fluid barrier that isolates the first reactant from the second reactant. Thereby, both of at least two reactant gases moving through the channel means can pass completely through the regenerative bed, quench the regenerative bed, absorb heat into the reaction gas, and then be combined to react with each other in the combustion zone.

[0071] In various embodiments, gases 30 and 32 (including fluids) can each contain components that react with components of the other reactants 30 and 32, and when combined, produce an exothermic reaction. For example, each of the first and second reactants can include one fuel gas and one oxidant gas that combust when combined with the other of the fuel and oxidant. By substantially separating the reactants, the location of the heat generated by the exothermic reaction can be controlled. In some embodiments, "substantially separated" can be defined to mean that at least 50 percent or at least 75 percent or at least 90 percent of the reactants having a minimal or limited stoichiometrically reactive amount of reactants between the first and second reactant streams are not consumed by the reaction until such time as these gases complete their axial movement through the recuperator 27. Thus, the majority of the first reactant 30 can be isolated from the majority of the second reactant 32, and the majority of the heat generated by the reaction combining the reactants 30 and 32 can occur after the reactants begin to exit the recuperator 27. The reactants can be gases, but optionally, some reactants can include liquids, mixtures, or gas phases.

[0072] The proportion of the reaction for these regeneration streams means the proportion of the reaction that is possible based on the stoichiometry of the overall feed. 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) to produce 10 volumes of H2O. In this case, if 10 volumes of hydrogen actually combust in the recuperation region (27), this represents a 100% reaction of the regeneration stream. This is because, in this example, the unreacted oxygen is present in an amount greater than the stoichiometric requirement, despite the presence of the remaining unreacted oxygen. Thus, in this example, hydrogen is the stoichiometrically limiting component. Using this definition, a reaction of less than 50% or less than 25% or less than 10% of the regeneration stream can occur during the axial movement of the recuperator (27).

[0073] In various embodiments, channels 28 and 33 can include a ceramic (including zirconium oxide), alumina, or other refractory material that can withstand temperatures exceeding 1200 °C, or 1400 °C, or 1600 °C. Additionally, or alternatively, channels 28 and 33 can have a wet surface area of -1 50 feet -1 to 3000 feet -1 or 100 feet -1 to 2500 feet -1 or 200 feet -1 to 2000 feet.

[0074] Referring again briefly to FIG. 2, 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. 2 and 3 are merely illustrative examples provided for illustrative purposes only and are not intended to represent comprehensive embodiments. The description of a "end" of a reactor merely means the end portion of the reactor with respect to the midpoint of the axis of the reactor. Thus, the description of a gas entering or exiting a "end" of a reactor, such as end 9, merely means that the gas enters or exits substantially at any of various points along the axis between the respective end surface of the reactor and the midpoint of the reactor, 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 surfaces, while the other is supplied to the respective ends of the reactor through slots or ports in the outer surface of the circumference or periphery on each end of the reactor.

[0075] Additional embodiments Embodiment 1. A method of operating a countercurrent reactor, comprising: in a reheating zone of the countercurrent reactor, exposing at least a portion of a fuel mixture containing fuel and at least 0.1% by volume of O2 to at least one heated surface to heat at least a portion of the fuel mixture; reacting the fuel mixture under combustion conditions in the reheating zone to form combustion gas; in a reaction zone of the countercurrent reactor, heating one or more regenerative surfaces to a regenerative surface temperature, the reaction zone containing a catalyst composition for an endothermic reaction and the reheating zone containing a catalyst composition for a supplementary exothermic reaction; exposing a reactant stream in the reaction zone to one or more regenerative surfaces to increase the temperature of the reactant stream; in the reaction zone, exposing the reactant stream to a catalyst composition for an endothermic reaction under endothermic reaction conditions to form a product stream, wherein the flow direction of the reactant stream in the reaction zone is opposite to the flow direction of the fuel mixture; in the reheating zone, exposing the product stream to a catalyst composition for a supplementary exothermic reaction under supplementary exothermic reaction conditions to form a reacted product stream; and in the reheating zone, heating at least one heated surface.

[0076] Embodiment 2. The method of Embodiment 1, wherein the catalyst composition for the endothermic reaction comprises a reforming catalyst.

[0077] Embodiment 3. The method of Embodiment 2, wherein the reactant stream contains a reformable hydrocarbon and steam, the product stream contains hydrogen, and the reformable hydrocarbon optionally contains methane.

[0078] Embodiment 4. The method according to any of the above embodiments, wherein the catalyst composition for the supplementary exothermic reaction comprises a water gas shift catalyst.

[0079] Embodiment 5. The method of Embodiment 4, wherein the water gas shift catalyst comprises a high temperature water gas shift catalyst, a second reforming catalyst optionally different from the reforming catalyst in the reaction zone, or a combination thereof.

[0080] Embodiment 6. The method according to any one of Embodiments 1 to 3, wherein the catalyst composition for the supplementary exothermic reaction comprises a methanol synthesis catalyst, a methanol conversion catalyst, or a combination thereof.

[0081] Embodiment 7. The method according to any one of Embodiments 1 to 3, wherein the catalyst composition for the supplementary exothermic reaction comprises an olefin oligomerization catalyst.

[0082] Embodiment 8. The method according to any one of Embodiments 1 to 3, wherein the catalyst composition for the supplementary exothermic reaction comprises a Fischer-Tropsch catalyst for olefin production.

[0083] Embodiment 9. The method according to any one of the above embodiments, wherein the fuel mixture burns downstream of the catalyst composition for the supplementary exothermic reaction with respect to the direction of flow of the fuel mixture.

[0084] Embodiment 10. The method according to any one of the above embodiments, wherein at least one of the fuel and O2 is not exposed to the catalyst composition for the supplementary exothermic reaction.

[0085] Embodiment 11. The method according to any one of the above embodiments, wherein the product stream is exposed to the catalyst composition for the supplementary exothermic reaction before passing through the external heat exchange stage.

[0086] Embodiment 12. The method according to any one of the above embodiments, wherein the combustion gas comprises a temperature of 420 °C or higher.

[0087] Embodiment 13. The method according to any one of the above embodiments, wherein the endothermic reaction conditions comprise a temperature of 600 °C or higher.

[0088] Although the present invention has been described and illustrated with reference to particular embodiments, those skilled in the art will recognize that the invention relates to variations not necessarily shown herein. For this reason, only the appended claims should be referred to in order to determine the true scope of the present invention.

Claims

Claim 1 A method of operating a reverse flow reactor, said method including a reforming step and a regeneration step including a supplementary exothermic reaction, wherein said regeneration step In the reheating zone of the reverse flow reactor, exposing at least a portion of the fuel mixture to at least one heated heat transfer surface to heat at least a portion of the fuel mixture, where the fuel mixture comprises fuel and at least 0.1% by volume of O ２ and; and comprises reacting said fuel mixture under combustion conditions in said reheating zone to form combustion gases, and passing said combustion gases through a reaction zone to heat one or more regeneration surfaces in a reaction zone of said reverse flow reactor; said reaction zone including one or more regeneration ceramic structures having a catalyst composition for an endothermic reaction on a surface of said one or more reaction zones; said reheating zone including one or more additional regeneration ceramic structures having a catalyst composition for a supplementary exothermic reaction; said at least one heated heat transfer surface being within at least one channel defined by one or more additional regeneration ceramic structures; wherein said reforming step comprises exposing a reactant stream in said reaction zone to a surface of said one or more reaction zones and increasing the temperature of said reactant stream; exposing said reactant stream to said catalyst composition for said endothermic reaction under endothermic reaction conditions in said reaction zone to form a product stream, wherein a flow direction of said reactant stream in said reaction zone is opposite to a direction of a flow of said fuel mixture; exposing said product stream to said catalyst composition for said supplementary exothermic reaction under supplementary exothermic reaction conditions in said reheating zone to form a reacted product stream, and heating at least one heat transfer surface in said reheating zone; wherein said catalyst composition for said supplementary exothermic reaction comprises any one of a) to c) below. a) A methanol synthesis catalyst, a methanol conversion catalyst or a combination thereof b) An olefin oligomerization catalyst c) A Fischer-Tropsch catalyst for olefin production Claim 2 The method according to claim 1, wherein said catalyst composition for said endothermic reaction comprises a reforming catalyst. Claim 3 The method according to claim 2, wherein said reactant stream comprises a reformable hydrocarbon and steam, and said product stream comprises hydrogen. Claim 4 The method according to any one of claims 1 to 3, wherein said catalyst composition for said supplementary exothermic reaction comprises a water gas shift catalyst. Claim 5 The method according to claim 4, wherein said water gas shift catalyst comprises a high temperature water gas shift catalyst. Claim 6 the catalyst composition for the supplementary exothermic reaction includes a water gas shift catalyst, the method according to claim 2, wherein the water gas shift catalyst includes a second reforming catalyst different from the reforming catalyst in the reaction region. **Claim 7** The method according to any one of claims 1 to 6, wherein the fuel mixture burns downstream of the catalyst composition for the supplementary exothermic reaction with respect to the direction of flow of the fuel mixture. **Claim 8** The fuel and the O ２ The method according to any one of claims 1 to 7, wherein at least one of them is not exposed to the catalyst composition for the supplementary exothermic reaction. **Claim 9** The method according to any one of claims 1 to 8, wherein the product stream is exposed to the catalyst composition for the supplementary exothermic reaction before passing through an external heat exchange stage. **Claim 10** The method according to any one of claims 1 to 9, wherein the combustion gas includes a temperature of 420 °C or higher. **Claim 11** The method according to any one of claims 1 to 10, wherein the endothermic reaction conditions include a temperature of 600 °C or higher.

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