Catalyst-sorbent structure for ammonia synthesis and sorbing, and method for producing ammonia.

Abstract

The active catalyst for ammonia synthesis is integrated with a special sorbent in the composition or composite such that the catalyst portion and the sorbent portion are in direct, close contact, thereby overcoming the thermodynamic limits of the conversion. The sorbent includes metal halide absorbents, zeolite adsorbents, and other material absorbents or adsorbents that can capture ammonia as it is produced in close or near molecular contact with the catalyst, and the composition / composite may be provided in the form of granular or pelletized structures. Essentially, by removing ammonia as it is formed, the forward reaction for ammonia production can be continued with virtually no attenuation, and as a result, a high net conversion rate can be achieved in a single pass cumulatively in a compartmentalized reactor operated in series.

Description

[Technical Field]

[0001] Cross-reference of related applications This PCT application claims the interests of U.S. Provisional Application No. 63 / 524,465, filed on 30 June 2023, and U.S. Provisional Application No. 63 / 506,744, filed on 7 June 2023, both of which are incorporated herein by reference in their entirety.

[0002] This disclosure relates, more broadly, to compositions and complexes for the synthesis of ammonia, and more specifically, to an operable catalyst-sorbent structure for ammonia synthesis, comprising a catalyst portion for ammonia synthesis that is in close contact with an sorbent portion for integrated ammonia sorbation during synthesis in a first reaction and sorption step, wherein the catalyst-sorbent structure is configured to provide synthesized ammonia in a subsequent desorption step, thereby providing ammonia in various forms, including in the form of a pellet structure for use in the ammonia synthesis, sorbation, and desorption steps. [Background technology]

[0003] Ammonia is an inorganic compound of nitrogen and hydrogen with the formula NH3. Ammonia is an extremely useful molecule essential for life. The discovery of the Haber-Bosch process in 1909, which catalytically converts nitrogen gas (N2) and hydrogen gas (H2) via an iron-based catalyst to produce ammonia, made it possible to synthesize and produce nitrogen fertilizers that support billions of lives. In 2019, approximately 88% of ammonia was used as fertilizer in the form of salts, solutions, or anhydrous materials. The use of ammonia as a fuel or fuel source is also increasing, and it is an effective fuel, particularly considering that it does not pollute the environment compared to carbon-based and renewable fuels, which contribute to global CO2 emissions.

[0004] Ammonia boils at -33.34°C (-28.012°F) at one atmospheric pressure, and at ambient temperature and pressure, ammonia is a colorless gas; therefore, liquid ammonia typically must be stored under pressure or at low temperatures. Because ammonia production is highly exothermic and limited by equilibrium, reactors are generally operated at high pressure (approximately 60–300 barg) and high temperature (approximately 400–600°C). Industrially, the conversion rate per pass from nitrogen and hydrogen gas feedstocks to ammonia production is low, although some processes, due to the thermodynamics of the Haber-Bosch process, have a conversion rate of 20% per pass. To achieve the high net utilization rate of feedstocks required for industrial processes, reactor effluents are cooled to the ammonia condensation temperature to capture the product and allow for the recycling of unreacted feedstocks. The recycled material is added to fresh feedstocks, heated and compressed, and fed back into the reactor inlet. Commercial processes, due to their low conversion rate per pass and high recycling ratio, require larger reactor volumes and have additional industrial energy requirements for heating / cooling and compression. The first-order reaction is shown in equation (1), and the enthalpy change (ΔH°) of the converted N2 is approximately -92 kJ / mol: N2 + 3H2 = 2NH3 (1)

[0005] A known process for synthesizing ammonia is disclosed in U.S. Patent No. 5,711,926, which discloses a Haber-Bosch ammonia synthesis reactor, followed by separate adsorption vessels for adsorbing the ammonia product. The separate devices are described as pressure swing adsorption (PSA) devices independent of the reactor. The disclosed system also includes a recycling line that returns unreacted nitrogen guiding gas and unreacted hydrogen feed gas to the upstream reactor.

[0006] Another process for synthesizing ammonia is described in U.S. Patents 9,914,645 and 10,287,173, which disclose a system comprising a first reactor having a catalyst bed, followed by a second reactor having an absorbent bed configured to selectively absorb at least a portion of the ammonia produced from the upstream reactor. The absorbent is disclosed to be regenerated by increasing the temperature or decreasing the pressure. However, this method requires stopping the flow of feed gas to the reactor and then decreasing the pressure in the container containing the absorbent or increasing the temperature. The disclosed system also includes a recycling line for unreacted nitrogen guiding gas, unreacted hydrogen feed gas, and unabsorbed ammonia that is fed back to the upstream reactor. Furthermore, the absorber system requires a larger temperature swing to carry out absorption at low temperatures of 50–100°C and desorption at 150–350°C.

[0007] Another process for synthesizing ammonia is disclosed in U.S. Patent No. 11,548,789, which discloses a system comprising a catalyst for ammonia production and an absorber configured to selectively absorb ammonia from the reaction mixture at a temperature of 180–330°C and a pressure of 1–20 bar. The disclosed system also includes a recycling line for unreacted nitrogen guiding gas, unreacted hydrogen feed gas, and unabsorbed ammonia that are fed back to the reactor.

[0008] U.S. Patent Publication 2020 / 0325030 discloses a catalyst and a metal halide absorbent as separate components arranged within the internal volume of a reactor. This process appears to utilize reaction temperatures above 350°C and employ a downstream separation unit.

[0009] Another process for synthesizing ammonia is disclosed in U.S. Patent No. 10,974,970, which has separate catalyst particles and absorbent particles, wherein the catalyst particles are at least 10 times larger than the sorbent particles and are introduced into a fluidized bed reactor, and the sorbent is circulated through multiple reactors as a continuous flow of solid sorbent particles.

[0010] Therefore, the ammonia synthesis industry needs compositions and processes that are more efficient and have higher net conversion rates. There is also a need to synthesize ammonia, as well as the generated ammonia, while minimizing the recycling of unreacted nitrogen and unreacted hydrogen feedstocks. Furthermore, there is a need for systems that can overcome the equilibrium thermodynamic limits for the conversion rate of hydrogen and nitrogen gas feedstocks to ammonia. In addition, there is a need for compositions that synthesize product ammonia gas from hydrogen and nitrogen, provide efficient adsorption via sorbents, and eliminate or reduce the need for condensers for separation by condensation. Cost-effective and environmentally friendly materials that enable ammonia synthesis at lower temperatures and / or lower pressures are also desired. [Overview of the project]

[0011] The inventors have remarkably discovered compositions, systems, and methods comprising multiple catalyst-sorbent particles that overcome the thermodynamic limitations conventionally encountered in converting nitrogen and hydrogen feedstock gases into ammonia gas. In some embodiments, the multiple catalyst-sorbent particles integrate an active catalyst for ammonia synthesis with a special sorbent for ammonia sorbent, which allows for the removal of ammonia at the time ammonia is essentially formed in the first reaction and sorbent step. The multiple catalyst-sorbent particles are also configured to provide an ammonia product obtained in a later desorption step, thereby releasing at least a portion of the ammonia synthesized and absorbed and / or adsorbed in the first reaction and sorbent step from the multiple catalyst-sorbent particles to provide the resulting ammonia product.

[0012] In some embodiments, the integration of the active catalyst and the special adsorbent includes a catalyst portion and an adsorbent portion configured within the same catalyst-adsorbent particle, and in some preferred embodiments, includes a plurality of catalyst-adsorbent particles, each having a catalyst portion and an adsorbent portion. In some embodiments, the integration of the active catalyst and the special adsorbent includes a close connection between the catalyst portion and the adsorbent portion, and in some embodiments, it is preferable that the catalyst portion and the adsorbent portion are in direct contact, and in some embodiments, it is more preferable that the catalyst portion is in molecular contact with or close to molecular contact with the adsorbent portion. In some embodiments, the plurality of catalyst-adsorbent particles can be provided in a compressed structure configuration such as compressed pellets, compressed tablets, granules, or extruded products. Alternatively, the pellets may be first formed from the active adsorbent material and then impregnated with the active catalyst material. This method avoids pressing the adsorbent and catalyst powder into the pellet simultaneously. The catalyst may be placed as an eggshell coating on the outside of the adsorbent pellet or may be placed throughout the entire adsorbent pellet. In some other alternative embodiments, the pellets may be formed from an active catalyst material or a supporting active catalyst material and then coated with an active sorbent material.

[0013] In some embodiments, instead of multiple particles, the integration of the active catalyst and the special adsorbent includes a monolithic structure. The integration of the active catalyst and the special adsorbent in the monolithic structure includes a close, direct connection between the catalyst portion and the adsorbent portion, preferably, in some embodiments, the catalyst portion is in molecular contact with or close to molecular contact with the adsorbent portion.

[0014] In some embodiments, the integration of the active catalyst and special adsorbent in a close, direct configuration between the catalyst and adsorbent portions, whether in the form of multiple catalyst-adsorbent particles or a monolithic structure, essentially allows for the removal of ammonia via absorption and / or adsorption as ammonia is formed through the catalytic reaction. By essentially removing ammonia as it is formed, the forward reaction for ammonia production can be continued with virtually no attenuation, and as a result, high net conversion rates can be achieved in a single pass or cumulatively within a reactor partitioned to operate in series, and the reaction rate of the catalytic reaction can be improved.

[0015] In some embodiments, the catalyst portion converts a non-condensable feedstock containing nitrogen and hydrogen into ammonia. The active catalyst material may include iron (Fe), cobalt (Co), ruthenium (Ru), molybdenum (Mo), or a combination thereof. The active catalyst material may be supported on a porous support material, e.g., a molecular, micro, or mesoporous support material, and it is generally understood that it may contain other promoters to increase catalytic activity and / or improve catalytic stability. In some other embodiments, the active catalyst material may be supported on an sorbent material such that the catalyst-sorbent is supported on a molecular porous support material and may contain other promoters to increase catalytic activity and / or improve catalytic stability. In some embodiments, the promoter may be potassium (K), cerium (Ce), cesium (Cs), barium (Ba), or a combination thereof.

[0016] In some embodiments, the molecular porous support material for the active catalyst or catalyst-adsorbent may have an average pore diameter of about 20 nm to about 50 microns, in some embodiments, about 50 nm to about 5 microns, and in some preferred embodiments, about 100 nm to about 1 micron.

[0017] In some embodiments, the molecular porous support material for the active catalyst or catalyst-adsorbent is approximately 1 to 1000 m2 It may have a surface area in the range of / gram.

[0018] In some embodiments, the porous support material may include alumina, silica, magnesium, ceria, titania, iron oxide, zeolite oxides, or combinations thereof. Other high-surface-area porous materials with reduced activity to the active material are also being considered.

[0019] In some other embodiments, the active catalyst may be self-supporting in a porous form. In some embodiments, a special adsorbent may be supported on the active catalyst.

[0020] The pores associated with the active catalyst or catalyst-sorbent may be linear or meandering, and facilitate gas-phase mass transfer through either molecular diffusion or Knudsen diffusion within them.

[0021] In some embodiments, the sorbent portion comprises one or more metal halide absorbents having an absorption affinity for NH3 than for N2 and H2. In some preferred embodiments, the sorbent portion comprises one or more metal halides, where the metal of the one or more metal halides is selected from Mn, Mg, Ca, and Fe, and the halide of the one or more metal halides is selected from Cl, Br, and Sr. In some other preferred embodiments, the sorbent portion comprises one or more zeolites, particularly microporous crystalline aluminosilicate materials such as Y-type zeolite, X-type zeolite, A-type zeolite, ZSM-5, or mixtures thereof.

[0022] In some embodiments, the sorbent portion may be a metal halide salt selected from the group consisting of LiCl, NH4Cl, CoCl2, MgCl2, CaCl2, MnCl2, FeCl2, NiCl2, CuCl2, ZnCl2, SrCl2, SnCl2, BaCl2, PbCl2, LiBr, NaBr, MgBr2, CaBr2, MnBr2, FeBr2, NiBr2, CoBr2, SrBr2, BaBr2, PbBr2, NH4Br, NaI, KI, CaI2, MnI2, FeI2, NiI2, SrI2, BaI2, NH4I, and PbI2. In some preferred embodiments, the sorbent portion may be a metal halide salt selected from the group consisting of MgCl2, CaCl2, MnCl2, FeCl2, and NiCl. In some other embodiments, one or more metal halide salts include MnCl2, MgCl2, CaCl2, MgBr2, CaBr2, MgClBr, CaClBr, MgCaBr, and mixtures thereof. In some preferred embodiments, the sorbent portion may include one or more zeolites, including zeolite Y, zeolite X, zeolite 13X, zeolite 4A, zeolite 5A, ZSM-5, or mixtures thereof. In some other preferred embodiments, the sorbent portion may be a mixture of one or more metal halide salts and one or more zeolites. Absorbents or adsorbents of other materials that capture ammonia when ammonia is generated in close contact or near molecular contact with the catalyst are also intended, and these may be used alone or in combination with one or more metal halide salts and / or one or more zeolites.

[0023] In some embodiments, the sorbent portion includes a metal halide MnCl2 that can absorb 6, 2, 1, 0.5, or 0 moles of ammonia per mole of metal halide, depending on the operating temperature. For example, when MnCl2 is used at an operating temperature of about 260°C to about 330°C, 1 mole of ammonia is absorbed per mole of absorbent. At an operating temperature of about 130°C to about 260°C, MnCl2 can absorb about 2 moles of ammonia per mole of absorbent. Below about 130°C, about 6 moles of ammonia can be absorbed per mole of MnCl2. At about 330°C to about 370°C, about 0.5 moles of ammonia can be absorbed per mole of MnCl2. Above about 370°C, ammonia absorption by MnCl2 is undesirable.

[0024] In some embodiments, the sorbing portion comprises one or more aluminosilicate zeolites, particularly one or more zeolites selected from zeolite Y, zeolite X, zeolite 4A, zeolite 5A, ZSM-5, or mixtures thereof. The one or more zeolites preferably bind to NH3 more than N2 and H2, and in some preferred embodiments, the affinity for NH3 is higher than the affinity for N2 and / or H2 by at least 5 times, at least 10 times, at least 100 times, at least 200 times, at least 300 times, at least 400 times, at least 500 times, at least 600 times, at least 700 times, and at least 1000 times or more.

[0025] In some embodiments, the sorbent is present in a temperature range of 100 to 500°C and a pressure range of 1 bar to 100 bar, in a concentration of 1 to 2000 mg. NH3 / g 収着剤 , or more preferably 5-300 mg NH3 / g 収着剤Ammonia can be absorbed within this volume range. In some embodiments, the adsorbent may be a material other than a metal halide, and may include, but is not limited to, a metal-organic framework (MOF), a covalent organic framework (COF), a zeolite imidazolate framework (ZIF), or a zeolite, or other adsorbent materials that selectively take up and adsorb NH3 in the gas phase within this temperature and pressure range. More preferably, the adsorbent takes up NH3 through a surface or bulk phase transition, where a sharp boundary exists between volumes under given conditions, similar to the exemplary MnCl2 material.

[0026] In some embodiments, the integration of an active catalyst and a special adsorbent includes a catalyst portion and an adsorbent portion that is mutually mixed and compressed into a configuration such as pellets, tablets, granules, or extruded articles. In some embodiments, the compressed catalyst-adsorbent configuration further includes a porous support material. In some embodiments, the active catalyst and the special adsorbent are mutually mixed such that the catalyst-adsorbent is configured on a monolithic structure.

[0027] In some other embodiments, the integration of the active catalyst and the special adsorbent includes an adsorbent portion configured to have a compressed structure such as a compressed pellet, compressed tablet, extruded product, or granules, which provides the adsorbent core, and the catalyst portion is coated on the adsorbent core such that the catalyst coating provides a surrounding shell or outer layer, thereby the catalyst portion at least partially encapsulates the adsorbent core, and preferably substantially encapsulates the adsorbent core. In some embodiments, the adsorbent core further comprises a porous support material. In some embodiments, the catalyst coating further comprises a porous support material. In some embodiments, the porous support material of the adsorbent core and the catalyst coating is the same material. In some embodiments, the porous support material of the adsorbent core and the catalyst coating is different material. In some embodiments, the catalyst coating, which comprises the active material and the support, has an average thickness of about 3 microns to about 200 microns, preferably about 10 microns to about 150 microns, more preferably about 20 microns to about 100 microns.

[0028] In some other embodiments, the integration of the active catalyst and the special adsorbent includes an adsorbent portion configured to have a compressed structure such as a compressed pellet, compressed tablet, or extruded product, which provides an adsorbent core; a catalyst portion is coated on the adsorbent core such that the catalyst coating provides a surrounding shell or outer layer, thereby the catalyst portion at least partially encapsulates the adsorbent core, preferably substantially encapsulates the adsorbent core; and a second adsorbent portion is coated on the catalyst coating as a surrounding shell or outer layer that at least encapsulates the catalyst coating, preferably substantially encapsulates the catalyst coating. In some embodiments, the adsorbent core further includes a porous support material. In some embodiments, the catalyst coating further includes a porous support material. In some embodiments, the adsorbent coating on the catalyst coating further includes a porous support material. In some embodiments, the adsorbent core, the catalyst coating, and the porous support material of the adsorbent coating are the same material. In some embodiments, the adsorbent core, the catalyst coating, and at least two of the porous support materials of the adsorbent coating are the same material. In some embodiments, the sorbent core, the catalyst coating, and the porous support material of the sorbent coating are all made of different materials. In some embodiments, at least two of the porous support materials among the sorbent core, the catalyst coating, and the sorbent coating are made of different materials. In some embodiments, the porous support materials of the sorbent core and the catalyst coating are made of different materials. In some embodiments, the catalyst coating has an average thickness of about 3 microns to about 200 microns, preferably about 10 microns to about 150 microns, more preferably about 20 microns to about 100 microns. In some embodiments, the sorbent coating has an average thickness of about 3 microns to about 200 microns, preferably about 10 microns to about 150 microns, more preferably about 20 microns to about 100 microns.

[0029] In some other embodiments, the sorbent portion and the catalyst portion are supported and dispersed along the same porous support (sequentially or simultaneously, for example, by initial wetting impregnation, colloidal synthesis, or sol-gel method, or by other means) so that catalyst-sorbent particles are supported on the same porous support. In some other embodiments, the sorbent portion is supported and dispersed on a porous support different from that of the catalyst portion.

[0030] The catalyst-sorbent configuration can be provided in a compressed structure such as compressed pellets, compressed tablets, or extruded products, as a plurality of catalyst-sorbent particles having an intermixed configuration, an sorbent core having a catalyst coating, or an sorbent core having a first catalyst coating and a second coating.

[0031] In some embodiments, each of the catalyst-sorbent pellets, tablets, extruded articles, or similar configurations used during the normal operation of ammonia synthesis preferably has an average diameter defined by a non-spherical hydraulic diameter of 1 to 20 mm, preferably 3 to 10 mm, and more preferably 3 to 9 mm.

[0032] In some embodiments, the catalyst-sorbent particles containing the support material have a partial load of active sorbent of 5% to 90% by weight, preferably 10% to 75% by weight, and more preferably 20% to 50% by weight. In some embodiments, the catalyst-sorbent particles have a partial load of active catalyst of 0.01% to 30% by weight, preferably 0.25% to 10% by weight, and more preferably 0.5% to 5% by weight. In some preferred embodiments, the catalyst-sorbent particles have a partial load of catalyst of less than about 5% by weight.

[0033] In some embodiments, the catalyst-sorbent particles have a (catalyst:sorbent) weight ratio of the amount of active catalyst supported to the amount of active sorbent supported, which is about 1:1 to about 1:100, preferably about 1:1 to about 1:50, and more preferably about 1:1 to about 1:10.

[0034] In some embodiments, the predicted partial sorbent loading density is approximately 100 kg / m³. 3~2000 kg / m 3 in the range of, preferably about 300 kg / m 3 ~ about 1500 kg / m 3 in the range of, more preferably about 500 kg / m 3 ~ about 1200 kg / m 3 is in the range of.

[0035] In some embodiments, the predicted catalyst partial loading density is in the range of about 10 kg / m 3 ~2000 kg / m 3 and preferably in the range of about 100 kg / m 3 ~ about 1500 kg / m 3 and more preferably in the range of about 150 kg / m 3 ~ about 1200 kg / m 3 is in the range of.

[0036] In some embodiments where an increased cycle time is desired, the weight loading of the catalyst portion can be in the range of about 5 kg / m 3 ~ about 500 kg / m 3 and preferably in the range of about 5 kg / m 3 ~ about 400 kg / m 3 and more preferably in the range of about 5 kg / m 3 ~ about 250 kg / m 3 while the weight loading of the sorbent portion can be in the range of about 50 kg / m 3 ~1500 kg / m 3 and preferably in the range of about 150 kg / m 3 ~ about 1400 kg / m 3 and more preferably in the range of about 250 kg / m 3 ~ about 1200 kg / m 3 is in the range of. In some embodiments, the (catalyst:sorbent) weight ratio of the catalyst portion loading to the sorbent portion loading can be in the range of about 1:3 to about 1:300.

[0037] In some embodiments, the integrated catalyst-sorbent process exceeds the equilibrium NH3 composition obtained without the sorbent portion. In some embodiments, the nitrogen conversion rate (in terms of the percentage of stoichiometric conversion of unreacted nitrogen feedstock to ammonia) may be in the range of about 30 to 99.99%, preferably 50 to 99.9%, and more preferably about 70 to 99%, per pass, and a pass is understood to be a single tube or a series of segmented tubes that are fluidly connected during one process feed cycle. In some embodiments, the hydrogen conversion rate (in terms of the percentage of stoichiometric conversion of unreacted hydrogen feedstock to ammonia) may be greater than 70% per pass, in some embodiments at least 80% to a maximum of 100%, in some other embodiments at least 80% to a maximum of 99.99%, and in some other embodiments at least 80% to a maximum of 99%, and a pass is understood to be a single tube or a series of segmented tubes that are fluidly connected during one process feed cycle.

[0038] In some embodiments, the integrated catalyst-sorbent of the present disclosure comprises a catalyst portion and an sorbent portion, wherein the catalyst portion can convert unreacted hydrogen feedstock and unreacted nitrogen feedstock into ammonia products, and the sorbent portion can absorb the generated ammonia; both the conversion of the catalyst portion and the absorption of the sorbent portion can occur at temperatures in the range of about 100°C to about 500°C, preferably about 200°C to about 400°C, more preferably about 250°C to about 350°C, and even more preferably about 280°C to about 330°C; both the conversion of the catalyst portion and the absorption of the sorbent portion can occur at pressures in the range of about 2 bar to about 200 bar, preferably about 5 bar to about 100 bar, more preferably about 5 bar to about 50 bar, and even more preferably about 5 bar to about 20 bar.

[0039] In some embodiments, the present disclosure relates to a process for producing ammonia, the process comprising providing the catalyst-sorbent of the present disclosure as a fixed bed, preferably a packed bed, in a reactor, under normal operating conditions, the catalyst portion converts unreacted hydrogen and unreacted nitrogen feedstocks into ammonia products, and the sorbent portion absorbs and / or adsorbs the generated ammonia.

[0040] In some embodiments, the catalyst-sorbent is placed in the reactor and is supported in an area of ​​about 0.1% to about 99.9% of the reactor volume, preferably about 10% to about 98%, preferably about 15% to about 96%, more preferably about 20% to about 94%, and in some embodiments, even more preferably about 25% to about 90% of the reactor volume.

[0041] In some preferred embodiments, the amount of catalyst-sorbent loaded in the reactor is at least 10% of the reactor volume, in some embodiments at least 20%, in some embodiments at least 30%, in some embodiments at least 40%, in some embodiments at least 50%, in some embodiments at least 60%, in some embodiments less than 98%, in some embodiments less than 80%, and in some embodiments less than 70%.

[0042] In some embodiments, the active catalytic portion of the catalyst-sorbent is present in the reactor in a weight range (w / w) of about 0.01% to about 20%, preferably about 0.25% to about 10%, and more preferably about 0.5% to less than 5%. In some preferred embodiments, the catalyst-sorbent particles have a catalyst portion load in the reactor of less than 5% by weight.

[0043] In some embodiments, the sorbent portion of the catalyst-sorbent is present in the reactor in a weight range (w / w) of about 5% to 95%, preferably 10% to 90%, and more preferably 20% to 80%.

[0044] In some embodiments, the catalyst-sorbent is present in the reactor in a catalyst-sorbent weight ratio of about 1:1 to about 1:300, preferably about 1:10 to about 1:50, and more preferably about 1:15 to about 1:25.

[0045] In some embodiments, a process for producing ammonia using the catalyst-sorbent of the present disclosure has a process cycle that is less than the complete sorbent capacity of the sorbent portion. In some embodiments, the process cycle is at least 20% to about 95% of the complete theoretical capacity, as defined by the operating temperature and operating pressure of the process bed.

[0046] In some embodiments, a process for producing ammonia using the catalyst-sorbent of the present disclosure comprises an initial process cycle having an initial conversion rate and a second process cycle having a second conversion rate, the second conversion rate having a lower conversion rate than the initial conversion rate, in some embodiments at least 0.1% lower than the initial conversion rate, and in some preferred embodiments 0.1% to 10% lower than the initial conversion rate.

[0047] In some embodiments, a process for producing ammonia includes providing the catalyst-sorbent of the present disclosure to multiple beds. In some embodiments, the multiple beds are provided in series. In some embodiments, the multiple beds are provided in parallel. In some embodiments, the multiple beds are provided in both series and parallel.

[0048] In some embodiments, unreacted hydrogen is supplied from a hydrogen source. It is intended that unreacted hydrogen can be supplied from any hydrogen source, but in some preferred embodiments, the hydrogen source includes synthesis from water in an electrolytic cell.

[0049] In some embodiments, unreacted nitrogen is supplied from a nitrogen source. While it is intended that unreacted nitrogen can be supplied from any nitrogen source, in some preferred embodiments the nitrogen source is a pressure swing adsorption (PSA) system, an air separation unit (ASU) system, a membrane separator, or a combination thereof.

[0050] The above summary is not intended to describe each of the exemplary embodiments or all implementations or aspects of the subject matter herein. The following drawings and descriptions illustrate more specifically various embodiments and aspects of this disclosure.

[0051] The subject matter of this specification can be better understood by considering the following detailed descriptions of various embodiments in relation to the accompanying drawings. [Brief explanation of the drawing]

[0052] [Figure 1] This is an exemplary depiction of the integration of an active catalyst and a special adsorbent in a mutually mixed catalyst-adsorbent particle configuration according to a particular embodiment of the present disclosure, thereby the catalyst portion and the adsorbent portion are mutually mixed with a support material and co-supported on the support material such that each catalyst-adsorbent particle contains a support particle, therein the exploded square diagram depicts a porous mutually mixed structure. [Figure 2] This is another exemplary description of the integration of an active catalyst and a special adsorbent in a catalyst-adsorbent particle configuration according to certain embodiments of the present disclosure, wherein the adsorbent portion may be a continuous or discontinuous layer on the surface of a support material, and the catalyst portion may be dispersed on the adsorbent surface and / or the surface of the support material to provide a co-supported catalyst-adsorbent configuration. [Figure 3] This is an exemplary depiction of the integration of an active catalyst and a special adsorbent in a compressed structural configuration such as a pellet, tablet, or extruded product, according to a particular embodiment of the present disclosure, thereby mixing the catalyst portion and the adsorbent portion in the compressed structural configuration, where the exploded ellipse diagram depicts a porous mixed structure. [Figure 4]This is an exemplary depiction of the integration of an active catalyst and a special adsorbent in a compressed structure such as a pellet, tablet, or extruded product according to a particular embodiment of the present disclosure, wherein the core of the adsorbent portion is at least partially enclosed by the coating of the catalyst portion, and thereafter, an exploded ellipse of the core depicts the porous structure of the adsorbent portion, and an exploded view of the interface between the core and the coating depicts the direct contact between the catalyst portion and the adsorbent portion. [Figure 5] This is an exemplary depiction of the integration of an active catalyst and a special adsorbent in a compressed structure such as a pellet, tablet, or extruded product according to certain embodiments of the present disclosure, wherein the core of the adsorbent portion is at least partially enclosed by the coating of the catalyst portion, and thereby the coating of the catalyst portion is at least partially enclosed by the coating of the adsorbent portion, where an exploded ellipse of the core depicts the porous structure of the adsorbent portion, and an exploded view of the interface between the core and the coating depicts the direct contact between the catalyst portion and the adsorbent portion. [Figure 6] This is a graphical representation of a particular embodiment of the present disclosure, showing the rate of ammonia production per hour per gram of catalyst as reported in the prior art as a filled circle, and the reaction rate data modeled for the comparative system of the present disclosure as an open square. [Figure 7] This includes a graphical depiction of the nitrogen gas conversion rate with temperature variation in a comparative analysis of prior art reported datasets with low catalyst load (0.2 g) and high catalyst load (1.5 g) according to a particular embodiment of the present disclosure, with modeled reaction rate data for a particular embodiment of a commercially available catalyst containing iron and cobalt, thereby reflecting a reasonable approximation of current and projected performance for scaling the catalyst portion. [Figure 8]A graph illustrating the modeled adsorption constants of ammonia (NH3) (depicted as a filled circle), nitrogen gas (N2) (depicted as a filled triangle), and hydrogen gas (H2) (depicted as an open circle) as a function of temperature over a range of test data provided in a particular embodiment of the present disclosure, thereby showing that the adsorption constant of H2 decreases with temperature, while the adsorption constants of NH3 and N2 increase with temperature. [Figure 9] A specific embodiment of this disclosure provides a graphical representation of the modeled floor temperature versus length of an integrated catalyst-sorbent based on catalyst loadings of 100-300 kg / m3 (100 kg / m3 is represented as a filled square, 125 kg / m3 as a filled triangle, 150 kg / m3 as a filled circle, 175 kg / m3 as an open X symbol, 200 kg / m3 as a filled triangle, and 300 kg / m3 as an open rhombus), thereby depicting higher catalytic activity or loading modeling data towards the front of the floor. [Figure 10] A graph illustrating the modeled bed temperature versus length of an integrated catalyst-sorbent based on catalyst loadings of 400 to 1200 kg / m3, according to a particular embodiment of the present disclosure (400 kg / m3 is depicted as a filled triangle, 800 kg / m3 as an open triangle, 1000 kg / m3 as a filled circle, and 1200 kg / m3 as an open circle), thereby illustrating that higher catalyst loadings have peak temperatures exceeding approximately 330°C. [Figure 11] A schematic representation of a fluid chamber having a void configuration including a packed bed in an annular chamber located between an inner porous tube and an outer heat transfer wall, according to a particular embodiment of the present disclosure, wherein process feeds during reaction and / or adsorption cycles are supplied to the annular packed bed, and an outlet of a central void space is closed, thereby allowing desorbed ammonia during a desorption cycle to exit the outlet of the central void space. [Figure 12A]This is a cross-sectional SEM image of a catalyst-adsorbent structure in pellet form according to a particular embodiment of the present disclosure. [Figure 12B] This is a cross-sectional EDS image of a catalyst-adsorbent structure in pellet form according to a particular embodiment of the present disclosure. [Figure 13] This is a graph of experimental data showing the weight increase of each ion-exchange zeolite due to ammonia adsorption according to a specific embodiment of the present disclosure. [Figure 14] This is a graph of experimental data illustrating the adsorption of ammonia by zeolite over time during a reaction process according to a particular embodiment of the present disclosure. [Figure 15] This graph shows experimental data illustrating the effect of pre-reduction on catalytic activity related to ammonia formation in a specific embodiment of the present disclosure. [Figure 16] This is a graph of experimental data illustrating the effect of ammonia pretreatment related to ammonia formation according to a specific embodiment of the present disclosure.

[0053] Various embodiments are possible, along with various modifications and alternative forms, the details of which are illustrated in the drawings and described in detail below. However, it should be understood that the intent is not to limit the claimed invention to the specific embodiments described. Rather, the intent is to encompass all modifications, equivalents, and alternatives that fall within the spirit and scope of the subject matter defined by the claims. [Modes for carrying out the invention]

[0054] As used herein, the terms “absorbent” or “solid absorbent” refer to and encompass salts of metal halide salts, metal-organic skeletons, and similar materials, thereby allowing ammonia to be present not merely on the surface of molecules, or on the surface of molecules that may form cage-like structures, or in addition, within the bulk material.

[0055] As used herein, the term “absorption” refers to the process by which a fluid (gas or liquid) enters the bulk phase of a solid material.

[0056] As used herein, the terms “adsorbent” or “solid adsorbent” refer to and include zeolites such as aluminosilicate zeolites, or other materials in which ammonia is present on the surface of molecules of a solid material, or on the surface of molecules of a solid material that forms a cage-like structure involving chemiadsorption, physiadsorption, or a combination thereof, rather than entering the bulk phase of the solid adsorbent material.

[0057] As used herein, the term “adsorption” refers to a process in which a fluid (gas or liquid) is retained on the surface of a solid material, including chemisorption.

[0058] As used herein, the term “adsorbent” refers to an insoluble material or mixture of materials used to recover a fluid through process absorption, adsorption, or both, and includes the terms absorbent, solid absorbent, adsorbent, and solid adsorbent.

[0059] As used herein, the term “intimate contact” refers to the relationship between the catalyst portion and the sorbent portion of a catalyst-sorbent structure with respect to characteristic dimensions less than the radius of an apparatus containing the catalyst and sorbent portions, e.g., particles, pellets, tablets, or extruded products, such that the catalyst and sorbent portions are contained within the same structural component.

[0060] As used herein, the term “coordination number” refers to the number of moles of ammonia retained per mole of sorbent portion.

[0061] As used herein, the term “capacity” refers to the weight in grams of ammonia held per gram of sorbent portion.

[0062] As used herein, the term “mutually mixed structure” refers to a granular catalyst-sorbent structure, such as particles, pellets, tablets, or extruded products, that contains a mutual mixture of catalyst and sorbent portions, including a homogeneous mixture of catalyst and sorbent portions.

[0063] As used herein, the term “single-coating structure” refers to a granulation catalyst-adsorbent structure, such as particles, pellets, tablets, or extruded products, that contains an inner core of one material and an outer shell coating comprising a second different material.

[0064] As used herein, the term “double-coated structure” refers to a granulation catalyst-sorbent structure, such as particles, pellets, tablets, or extruded products, comprising an inner core of one material, a first outer shell coating containing a second different material, and a second outer shell coating containing a third material. The third material may be either the first or second material, a combination of the first and second materials, the first or second material having different weight loads of the active material, or any combination thereof.

[0065] As used herein, the term “continuous-discontinuous sorbent” refers to a region of a granulated catalyst-sorbent structure that is continuous with respect to a first portion of the structural radius and discontinuous with respect to a second portion of the structural radius. The length of the continuous region may range from approximately 0.05 to 5 times the structural radius. The length may be greater than the structural radius, according to the nonlinear and meandering pore structure that may be found within the catalyst-sorbent structure.

[0066] The term "continuous-discontinuous" may refer to areas of a catalyst or adsorbent coating that are continuous to a portion of the surface but discontinuous to a portion of the surface coating of particles, pellets, tablets, or extruded products.

[0067] As used herein, the term “continuous sorbent” refers to a continuous and molecularly connected sorbent moiety dispersed across a porous catalyst-sorbent structure, thereby ensuring that at least 95% of the sorbent molecules in the sorbent moiety are in molecular contact.

[0068] As used herein, the term “discontinuous sorbent” refers to an sorbent portion of a catalyst-sorbent in which less than 5% of the sorbent molecules are in molecular contact, resulting in the sorbent portion being dispersed throughout as islands, points, or small connections of molecules separated from other islands or discontinuous regions of the sorbent. Molecules of a support, catalyst, or other additive that allow gaseous diffusion, and open pores, may separate discontinuous small groups of absorbent molecules.

[0069] As used herein, the term “dual function” refers to a material that provides two functions, such as a material that functions as a catalytic portion to increase the reaction rate or synthesis rate of ammonia, and a material that also functions as an sorbent portion or a material that activates an sorbent portion to increase the volume of ammonia during a combined reaction-adsorption operation.

[0070] As used herein, the terms “co-extruded catalyst-sorbent” or “co-extruded catalyst-sorbent structure” refer to the mixing of powders or particles containing catalyst and sorbent portions before being compressed into a structure such as a pellet, tablet, or extruded product, the resulting catalyst-sorbent structure retains internal porosity to and from active sites for the reaction, adsorption, or desorption of reactants and products, and combinations thereof. Extruded structures, including pellets, tablets, and ore extruded products, may contain binders used to assist material processing during synthesis, thereby it is understood that such binders are removed after extrusion by thermal processes, extraction, or other means that substantially remove the binder material from the finished structure so that the pellet retains internal porosity to accommodate internal mass diffusion of reactants and products.

[0071] As used herein, the term “co-supported catalyst-adsorbent” refers to a structure in which the catalyst and adsorbent are supported on the same support particles. If the catalyst is self-supported, the support may be the catalyst itself.

[0072] As used herein, the terms “catalyst loading density” and “partial catalyst loading density” refer to the mass (kg) of catalyst loaded or packed within the volume of a reaction tube, chamber, or vessel, thereby kg / m³ 3 The catalyst load represents the amount of catalyst in the reaction chamber. The catalyst mass includes the active catalyst material, any supporting material for the catalyst, and any additional dopants or additives added to improve the stability or activity of the catalyst system.

[0073] As used herein, the terms “sorbent loading density” and “partial sorbent loading density” refer to the mass (kg) of sorbent loaded or filled within the volume of a reaction tube, chamber, or container, thereby kg / m³ 3 The amount of sorbent loaded represents the amount of sorbent in the reaction chamber. The mass of the sorbent includes the active sorbent material, any supporting material for the sorbent, and any additional dopants or additives added to improve the stability or activity of the sorbent portion.

[0074] As used herein, the term “cycle time” means the time (in units of hours) during which a feed of unreacted hydrogen and / or unreacted nitrogen is continuously supplied to a first process vessel containing the catalyst-sorbent structure of the present disclosure, the feed of unreacted hydrogen and / or unreacted nitrogen being converted, either entirely or partially, to ammonia. During the same cycle time, a second vessel containing another catalyst-sorbent structure may be regenerated, thereby removing, either entirely or partially, the ammonia sorbed during the previous cycle time.

[0075] As used herein, the term “feed switching” refers to the end of a cycle time in which the feed of unreacted hydrogen and / or unreacted nitrogen is moved to at least a second process vessel containing other catalyst-sorbent structures, thereby allowing the reaction to continue forming ammonia. The feed of unreacted hydrogen and / or unreacted nitrogen is moved in parallel from one or more vessels to at least a second vessel or a set of parallel vessels by using valves to block the flow to one part of the system while opening the flow to at least a second part of the system.

[0076] As used herein, the terms “partitioned reactors and absorbers” and “partitioned beds” refer to a series of reactor and absorber tubes that are fluidly connected during each portion of the cycle time. The dimensions, operating conditions, and loads of the absorbent, and / or the loads of the catalyst, may differ or be the same in the series-connected partitioned tubes.

[0077] As used herein, the term “recycle reactor” refers to a reactor system comprising a first vessel containing a catalyst for producing ammonia from nitrogen and hydrogen, followed by at least a second unit operation for removing all or part of the ammonia from the product effluent leaving the first vessel. The unreacted feed of unreacted hydrogen and / or unreacted nitrogen is recompressed and reheated to inlet conditions for the first vessel to continue the reaction.

[0078] As used herein, the term “recycle ratio” refers to a system expressed as the amount of product selectively separated downstream of the first reactor divided by the conversion rate per pass. Unreacted feedstock of unreacted hydrogen and / or unreacted nitrogen after product separation by absorption, adsorption, membrane, concentration, or any other method is recycled to the reactor inlet. The recycled feedstock of unreacted hydrogen and / or unreacted nitrogen is recompressed to the inlet pressure and heated to the inlet temperature.

[0079] As used herein, the term “redox activation” refers to one or more process steps performed in a hydrogen-containing gas, including an oxidation step followed by a reduction step, prior to operating a catalyst for ammonia synthesis. In one embodiment, a series of redox steps are performed such that the catalyst can be oxidized at least twice in an intermediate reduction step and a final reduction step before operating the catalyst for ammonia synthesis. The catalyst portion of this disclosure is operated in a reduced state. In an alternative embodiment, the catalyst undergoes all steps of redox activation before being supported or packed into the reactor tube, so that only the final reduction step is performed in situ before operation.

[0080] As used herein, the term “activation” refers to a process step prior to the operation of the catalyst-sorbent for the production and sorbing of ammonia, thereby increasing the activity, capacity, or stability of the catalyst and / or sorbent portion. The activation step can be carried out at a temperature and / or pressure that may be higher, lower, or the same as the operating conditions of the ammonia synthesis reaction. The flow rate and gaseous composition of the fluid during activation may differ from those in the reactor and during the sorbing operation.

[0081] As used herein, the terms “absorbent transfer,” “adsorbent transfer,” or “adsorbent transfer” refer to the incorporation of ammonia into the sorbent portion of a catalyst-sorbent compound by a chemical reaction or the formation of a complex. For example, without limiting the terms “absorbent transfer” or “adsorbent transfer,” if the sorbent portion contains MnCl2, the incorporation may have the following complex compound: MnCl2-xNH3+yNH3→MnCl2-zNH3 (wherein x+y=z, and x and z may have values ​​of 0, 0.5, or integer values ​​such as approximately 1, 2, 4, 6, or 8).

[0082] As used herein, the term “transition temperature” refers to the temperature at which an absorber transition occurs, and is typically set by the pressure of NH3.

[0083] As used herein, the term “surface velocity” refers to the gas phase velocity at local temperature and pressure in the open cross-section of a reactor tube or vessel.

[0084] As used herein, the term “average residence time” refers to the average time that reactant molecules spend in a reactor system containing a catalyst-sorbent structure, which is calculated by dividing the surface velocity by the product of the reactor length and porosity.

[0085] As used herein, the term “metal halide” refers to a material comprising at least one metal molecule and at least one halide molecule that form a stable molecular complex having an affinity for ammonia. It is desirable that the metal halides of this disclosure can remove ammonia molecules from a fluid (such as a gas phase) and retain the ammonia molecules within a stationary phase containing the metal halide.

[0086] As used herein, the term "zeolite" refers to a microporous crystalline aluminosilicate material having an affinity for ammonia. It is desirable that the zeolite of this disclosure can remove ammonia molecules from a fluid (such as a gas phase) and retain ammonia molecules on a surface within a stationary phase containing the zeolite.

[0087] As used herein, the term “decomposition reaction” refers to the reverse reaction of ammonia synthesis, in which ammonia is catalyzed and decomposed into gaseous nitrogen and hydrogen. It is desirable to reduce the decomposition rate so that most of the ammonia formed during the synthesis step can be removed from the system to be recovered as a useful process product.

[0088] As used herein, the terms “pellets,” “tablets,” and “extruded products” refer to granular structured materials of catalyst-sorbents, thereby configured to flow substantially freely when filled or supported in a reactor vessel, and the granular catalyst-sorbent structures contain internal porosity to facilitate the diffusion of reactants and products to and from active sites for reaction, adsorption, desorption, and combinations thereof.

[0089] The term "tpd" refers to the tonnage of ammonia nominally formed from the catalyst-sorbent of this disclosure per day, where 1 ton is understood to be 1,000 kg of ammonia. While the nominal capacity of the plants disclosed herein relates to 4 tpd, it should be understood that the techniques of the present invention can be applied to systems of smaller or larger capacity by decreasing or increasing the number of parallel reactor tubes or vessels, respectively, and / or by increasing the length and / or diameter of the reactor to the extent that performance is possible based on heat transfer in the packed bed. In one embodiment, the plant capacity of the present invention may range from about 1 to 5,000 tpd or more, more preferably about 2 to 100 tpd of ammonia production.

[0090] The inventors have remarkably discovered compositions, systems, and methods that overcome the thermodynamic limitations conventionally encountered in converting nitrogen and hydrogen feedstock gases into ammonia gas. In some embodiments, by integrating an active catalyst for ammonia synthesis with a special sorbent for ammonia sorbent in a common structural component, it becomes possible to remove ammonia essentially during ammonia formation. Solvation can take the form of absorption, adsorption, or a combination thereof.

[0091] Catalyst-adsorbent particles and structure Referring here to Figures 1-5, the integration of the active catalyst and special adsorbent is disclosed at the microscale level at the individual catalyst-adsorbent particle level or at the multiple catalyst-adsorbent particle level, as well as in operable ammonia synthesis structures such as pellets, tables, and extruded products, which in some embodiments may consist of multiple catalyst-adsorbent particles, or in other embodiments may consist of different microscale structures.

[0092] Referring here to Figures 1 and 2, the integration of the active catalyst and the special adsorbent can take the form of individual catalyst-adsorbent particles 10. For example, as illustrated in Figure 1, the integration of the active catalyst and the special adsorbent can have the configuration of mutually mixed catalyst-adsorbent particles 10, thereby the catalyst portion 20 and the adsorbent portion 10 are mutually mixed with and co-supported on the support material 30, and as a result, each catalyst-adsorbent particle contains a support particle 30 having mutually mixed catalyst portion 20 and adsorbent portion 10. The catalyst portion 20 and the adsorbent portion 10 can be seen on the surface of the support material 30, such as within multiple pores 32 of the porous support material 30. As shown in Figure 2, the adsorbent portion 10 can take the form of a continuous / discontinuous layer on the support surface 30. The catalyst portion 20 is shown dispersed on the adsorbent portion 10, but the catalyst portion 20 can be dispersed on the adsorbent portion 10 and / or on the support surface 30. Figures 1 and 2 illustrate a co-supported catalyst-sorbent particle configuration according to a particular embodiment of the present disclosure.

[0093] Referring to Figures 3-5, various operational ammonia synthesis structures are shown. Referring here to Figure 3, multiple intermixed structures 100 are shown for producing gaseous products such as ammonia (NH3) from unreacted nitrogen and unreacted hydrogen feedstocks. The intermixed structure 200 is a granulated catalyst-sorbent structure having multiple catalyst-sorbent particles, each catalyst-sorbent particle having an intermixed sorbent portion 110 with a catalyst portion 120. In some preferred embodiments, the multiple catalyst-sorbent particles are compressed into the intermixed structure 100. The compressed intermixed structure 200 can be composed of particles, pellets, tablets, or extruded products. The sorbent portion 110 and catalyst portion 120 can be intermixed into a homogeneous or heterogeneous intermix. The exploded elliptic diagram of the intermixed structure 200 depicts porosity, which may be linear or meandering, and facilitates gas-phase mass transfer through either molecular diffusion or Knudsen diffusion within it.

[0094] Referring here to Figure 4, another operational ammonia synthesis structure is shown, which is a single-coat structure 200 for producing a gaseous product such as ammonia (NH3) from a feedstock of unreacted nitrogen and a feedstock of unreacted hydrogen. The single-coat structure 200 can be a granulated catalyst-sorbent structure having an inner core 205 containing an sorbent portion 210 and an outer shell coating 215 containing a catalyst portion 220. In some preferred embodiments, the sorbent portion 210 is compressed into a desired configuration so that the inner core 205 can be composed of particles, pellets, tablets, or extruded products. In some embodiments, the sorbent portion 210 may be supported on a support material such as a porous support material so that the inner core 205 contains the sorbent portion 210 and the support material. The outer shell coating 215 can form a continuous outer shell that essentially encloses the inner core 205. In some other embodiments, the outer shell coating 215 can form a discontinuous outer shell such that the outer shell coating 215 at least partially encloses the inner core 205. In other embodiments, the outer shell coating 215 may be a continuous-discontinuous structure in which part of the outer coating 215 is continuous around the inner core 205 structure and part is discontinuous. An exploded elliptic of the inner core 205 depicts the porous structure of the sorbent portion, where the porosity may be linear or meandering and facilitates gas-phase mass transfer through either molecular diffusion or Knudsen diffusion within it. The outer shell coating 215 may have a porous structure, where the porous structure may be linear or meandering and facilitates gas-phase mass transfer through either molecular diffusion or Knudsen diffusion within it. An exploded view of the interface between the inner core 205 and the outer shell coating 215 depicts the close contact between the sorbent portion 210 and the catalyst portion 220.

[0095] Referring here to Figure 5, another operational ammonia synthesis structure is shown, which is a double-coated structure 300 for producing a gaseous product such as ammonia (NH3) from a feedstock of unreacted nitrogen and a feedstock of unreacted hydrogen. The double-coated structure 300 is a granulated catalyst-sorbent structure having an inner core 305 containing an sorbent portion 310, an inner shell coating 315 containing a catalyst portion 320, and an outer shell coating 325 containing a second sorbent portion 330. In some preferred embodiments, the sorbent portion 310 is compressed into a desired configuration so that the inner core 305 can be composed of particles, pellets, tablets, or extruded products. In some embodiments, the sorbent portion 310 may be supported on a support material such as a porous support material so that the inner core 305 contains the sorbent portion 210 and the support material. The inner shell coating 315 can form a continuous inner shell that essentially encloses the inner core 305. The outer shell coating 325 can form a continuous or continuous-discontinuous outer shell that essentially encloses the inner shell coating 315. In some other embodiments, the inner shell coating 315 can form a discontinuous shell such that the inner shell coating 315 encloses the inner core 305 at least partially. Furthermore, the outer shell coating 325 can form a discontinuous shell such that the outer shell coating 325 encloses the inner shell coating 315 at least partially. The exploded elliptic diagram of the inner core 305 depicts the porous structure of the sorbent portion, where the porosity may be linear or meandering and facilitates gas-phase mass transfer through either molecular diffusion or Knudsen diffusion within it. The inner shell coating 315 and / or outer shell coating 325 can also have a porous structure, where the porous structure may be linear or meandering and facilitates gas-phase mass transfer through either molecular diffusion or Knudsen diffusion within it. The exploded view of the interface between the inner core 305, the inner shell coating 315, and the outer shell coating 325 illustrates the close contact between the catalyst portion 320 and the respective adsorbent portions 310 and 330.

[0096] Catalyst-adsorbent structures can have dual-function materials, where the same material can perform two functions. In some embodiments, catalyst-adsorbent structures having dual-function materials perform two functions, but the catalyst-adsorbent structure contains separate materials for catalytic conversion and sorbing that are maintained in close contact within the catalyst-adsorbent structure, thereby the maximum distance between the sorbent site and the catalyst site is less than the radius of the catalyst-adsorbent structure, such as a pellet.

[0097] The sorbent portion can be arranged discontinuously or in a continuous-discontinuous sorbent configuration within the catalyst-sorbent structure so that volume expansion is localized to maximize the mechanical stability of the catalyst-sorbent structure. This makes it possible to circulate the catalyst-sorbent structure for at least several hundred to tens of thousands of times without damage.

[0098] The catalyst-sorbent structures of the present disclosure, as illustrated in Figures 1-3, may comprise a plurality of particles compressed into an operable ammonia synthesis structure, such as pellets, tablets, granules, or extruded articles, each particle integrating the catalyst portion into close contact with the sorbent portion, thereby enabling the removal of ammonia as ammonia is formed via a catalytic reaction. Alternatively, as shown in Figures 4-5, the operable ammonia synthesis structure may have an inner core containing the sorbent portion, providing a different configuration in which the catalyst-sorbent structure integrates the catalyst portion into close contact with the sorbent portion, which also enables the removal of ammonia as ammonia is formed via a catalytic reaction. Alternatively, the catalyst-sorbent structure may be a monolithic structure such that the catalyst portion and the sorbent portion coexist on a common monolithic structure, which also enables the removal of ammonia as ammonia is formed via a catalytic reaction, and the catalyst portion and the sorbent portion are preferably configured to be in direct contact with the monolithic structure. The catalyst-sorbent structure of this disclosure can essentially continue the forward reaction for ammonia production with little attenuation by removing ammonia as ammonia is formed, and as a result, a high net conversion rate can be achieved in a single pass or cumulatively in a reactor partitioned to operate in series during the cycle time before feed switching.

[0099] The catalytic portion of the catalyst-sorbent structure of this disclosure is capable of converting a non-condensable feedstock containing unreacted nitrogen and unreacted hydrogen into ammonia, particularly gaseous ammonia. The active catalytic material of the catalytic portion may include iron, cobalt, ruthenium, molybdenum, or a combination thereof. The active catalytic material may be supported on a molecular, micro, or mesoporous support material and may contain other promoters to increase catalytic activity and / or improve catalytic stability. The active catalytic material may be supported on an sorbent material such that the catalyst-sorbent is supported on a molecular porous support material and may contain other promoters to increase catalytic activity and / or improve catalytic stability. Alternatively, the active catalyst may be self-supported in a porous form. In some embodiments, the sorbent may be supported on the active catalyst.

[0100] The support material for the catalyst portion and / or adsorbent portion may be an oxide material or other high surface area porous material. Exemplary oxide materials include alumina, silica, magnesium, ceria, titania, iron oxide, and combinations thereof. The support material is preferably a molecular porous support material having an average pore diameter of about 20 nm to about 50 microns, in some embodiments about 50 nm to about 5 microns, and in some preferred embodiments about 100 nm to about 1 micron. The molecular porous support material is preferably about 1 m 2 / gram ~ 1000m 2 It has a surface area in the range of / grams.

[0101] The pores associated with the active catalyst or catalyst-sorbent may be linear or meandering, and facilitate gas-phase mass transfer through either molecular diffusion or Knudsen diffusion within them.

[0102] In some preferred embodiments, the sorbent portion comprises one or more metal halide absorbents having an absorption affinity for NH3 than for unreacted nitrogen (N2) and unreacted hydrogen (H2). In some preferred embodiments, the sorbent portion comprises one or more metal halides, wherein the metal of the one or more metal halides is selected from Mn, Mg, Ca, and Fe, Sr, and the halide of the one or more metal halides is selected from Cl, Br.

[0103] The sorbent portion can be a metal halide salt selected from the group consisting of LiCl, NH4Cl, CoCl2, MgCl2, CaCl2, MnCl2, FeCl2, NiCl2, CuCl2, ZnCl2, SrCl2, SnCl2, BaCl2, PbCl2, NH4Cl, LiBr, NaBr, MgBr2, CaBr2, MnBr2, FeBr2, NiBr2, CoBr2, SrBr2, BaBr2, PbBr2, NH4Br, NaI, KI, CaI2, MnI2, FeI2, NiI2, SrI2, BaI2, NH4I, and PbI2. In some preferred embodiments, the sorbent portion is a metal halide salt selected from the group consisting of MgCl2, CaCl2, MnCl2, FeCl2, and NiCl. In some other embodiments, one or more metal halide salts include MnCl2, MgCl2, CaCl2, MgBr2, CaBr2, MgClBr, CaClBr, MgCaBr, and mixtures thereof. Absorbents or adsorbents of other materials that capture ammonia are also intended when ammonia is produced in close contact or near molecular contact with the catalyst.

[0104] One preferred sorbent portion comprises a metal halide MnCl2 that can absorb 6, 2, 1, 0.5, or 0 moles of ammonia per mole of metal halide, depending on the operating temperature. For example, when MnCl2 is used at an operating temperature of about 260°C to about 330°C, 1 mole of ammonia is absorbed per mole of absorbent. At an operating temperature of about 130°C to about 260°C, MnCl2 can absorb about 2 moles of ammonia per mole of absorbent. Below about 130°C, about 6 moles of ammonia can be absorbed per mole of MnCl2. Between about 330°C and about 370°C, about 0.5 moles of ammonia can be absorbed per mole of MnCl2. Above about 370°C, ammonia absorption by MnCl2 is undesirable.

[0105] In some embodiments, the sorbent is present in a temperature range of 100 to 500°C and a pressure range of 1 bar to 100 bar, in a concentration of 1 to 2000 mg. NH3 / g 収着剤 , or more preferably 5-300 mg NH3 / g 収着剤 It can absorb ammonia within this capacity range.

[0106] In some embodiments, the adsorbent may be a material other than a metal halide, and may include, but is not limited to, a metal-organic framework (MOF), a covalent organic framework (COF), a zeolite imidazolate framework, or a zeolite, or other adsorbent materials that selectively incorporate NH3 in the gas phase within this temperature and pressure range. More preferably, the adsorbent incorporates NH3 through a surface or bulk phase transition, where a sharp boundary exists between capacities under given conditions, similar to the exemplary MnCl2 material.

[0107] In some other preferred embodiments, the sorbent portion of the catalyst-sorbent structure comprises one or more zeolites having an adsorption affinity for NH3 than for unreacted nitrogen (N2) and unreacted hydrogen (H2) on the surface of the zeolite material. The affinity for NH3 on the surface of the zeolite material can be due to chemiadsorption, physiadsorption, size exclusion, or a combination thereof. Compared to metal halides having a step isotherm, zeolites typically have a smooth isotherm such that NH3 adsorption increases with NH3 pressure / concentration at a constant temperature. Similarly, zeolites have a smooth isobaric curve, where NH3 adsorption decreases as the temperature increases at a constant NH3 pressure. While any zeolite may be used as the adsorbent in the catalyst-sorbent particles of this disclosure, preferred zeolites include zeolite Y, zeolite X (particularly 13X), zeolite 4A, zeolite 5A, ZSM-5, or mixtures thereof. Each of the aforementioned zeolites may be in hydrogen form, sodium form, and / or contain other cations. Other cations in zeolites include transition metals, alkali metals, or rare earth metals such as Mg, Mn, Cu, Co, Ru, Fe, K, Ce, Cs, and Zn.

[0108] While it is intended that any zeolite may be used as the adsorbent in the catalyst-sorbent particles of this disclosure, there may be cases where one or more particular zeolites are preferred based on various parameters. For example, some of the main differences between zeolites stem from effective pore size and material hydrophobicity. If the pore size tends to be the same size as the gas moving through the sorbent, there may be steric exclusion in internal molecular transport, including internal pellet molecular transport. Since NH3 is a small molecule with a diameter of 3-4 Å, if the pore size of a zeolite material such as zeolite 4A has a pore diameter of approximately 4 Å, for example, molecules larger than ammonia are typically excluded from the pores based on size exclusion, while ammonia tends to pass through the pores and adsorb to oxygen ions or cations within the porous framework.

[0109] In some preferred embodiments, the zeolite can have a pore size smaller than N2 and H2 molecules but larger than NH3 molecules, resulting in a pore size that allows NH3 flow while excluding unreacted N2 and H2. In some embodiments, the pore size is about 3 Å to about 5 Å, more preferably about 4 Å to about 5 Å. When the zeolite has an effective pore size of less than about 5 Å, the pore size results in a molecular flow size exclusion that provides selectivity for NH3. The size exclusion of the zeolite pores allows NH3 to pass through the pores but excludes larger molecules, including unreacted nitrogen and hydrogen. Once inside the porous framework of the zeolite, NH3 tends to be absorbed by oxygen ions or cations.

[0110] In some other preferred embodiments, the zeolite has an effective pore size much larger than that of the NH3 molecule. In some embodiments, the effective pore size of the zeolite is greater than 5 Å, for example at 13X, thereby the selectivity for NH3 is achieved through the surface properties of the zeolite.

[0111] In some embodiments, the surface polarity of the zeolite and its pores can cause preferential binding of polar NH3 to H2 or N2, as measured by the selectivity ratio or sorption capacity ratio of NH3 to H2 or N2. In some embodiments, the sorption capacity of NH3 is at least 5 times, preferably 10 times or more, for NH3 relative to H2 or N2.

[0112] The pore size of commercially available known adsorbents containing zeolites can be manipulated to a desired pore size. Zeolites are typically found in Na or H form, where Na or H is the primary cation in the crystalline structure. Ion exchange may partially or completely replace the Na and / or H cations with other metals, such as alkalis or transition metals. By changing the metal in the zeolite, the pore size is affected, and the selectivity to NH3 may increase. It can also affect the binding strength to NH3, and therefore the NH3 adsorption capacity of the zeolite. In some embodiments, larger ions bind to NH3 more weakly. The substituted ions can be monovalent, divalent, or trivalent. The electric field inside and on the surface of the zeolite may change due to ion exchange, and the results of ion exchange can be changes in pore size, surface polarity, electric field gradient, and / or hydrophobicity. These changes may result in an increase or decrease in activation energy and / or the binding energy of NH3 to the zeolite, and thus can alter the NH3 capacity of the sorbent at a given temperature and pressure. It can be understood that it is desirable to increase the NH3 capacity at a given temperature and pressure, preferably 250-400°C, more preferably 300-350°C. It is also desirable to increase the NH3 capacity of the zeolite at a given temperature / pressure pair compared to another temperature / pressure pair, thereby allowing the separation of ammonia from the gas stream via temperature / pressure swings.

[0113] In some embodiments, the sorbent is a zeolite selected from Y, X, 4A, 5A, ZSM-5, or others that preferentially binds to NH3 rather than unreacted N2 and unreacted H2. The sorbent may comprise one or more of the aforementioned zeolites. In some embodiments, the sorbent composition may vary radially or axially throughout the reactor bed. For example, certain sorbent compositions may be preferred toward the reactor inlet end where the reactant gas is introduced, rather than toward the reactor outlet end (where NH3 and unreacted H2 and unreacted N2 are discharged from the reactor). In another exemplary embodiment, it may be preferable to have one type of sorbent composition toward the center of the reactor bed rather than toward the ends of the reactor bed where a different sorbent composition may be desired.

[0114] In some embodiments, the zeolite can act as both an adsorbent and a support. The zeolite can be impregnated with or otherwise supported by an active catalyst or secondary adsorbent. In some preferred embodiments, the secondary adsorbent is another zeolite material. In some other preferred embodiments, the secondary adsorbent is a metal halide. The secondary adsorbent may be a metal halide in the form of a coating or cluster within the pores of the zeolite.

[0115] In some embodiments, the active catalyst for NH3 synthesis using a zeolite sorbent may be Fe, Co, Ru, Mo, or a combination thereof. The active catalyst may be supported in the macropores or micropores of the zeolite, and may take the form of nanocrystals, clusters, coatings, or similar forms. The active catalyst may also be in the form of metal ions. For example, the Na ions in zeolite 13X can be replaced with an active ammonia catalyst such as Ru and / or Co.

[0116] In some preferred embodiments, the zeolite may be impregnated, ion-exchanged, coated, or otherwise supported with a promoter material. The promoter material is preferably K, Ce, Cs, Ba, or a mixture thereof. The promoter material is preferably supported on the zeolite in an amount greater than 0 to a maximum of about 10% by weight. In exemplary embodiments, it may be beneficial to incorporate Ce into the framework of the zeolite material to improve catalytic activity, for example, using a Ru catalyst.

[0117] In some embodiments, it may be preferable to reduce the oxide layer in relation to the active catalyst in order to achieve maximum activity. The active catalyst can be reduced by applying a temperature above 400°C to the catalyst-sorbent particles (e.g., about 400°C to about 500°C for Fe as the active catalyst). Reduction of the catalyst at temperatures of 400 to 500°C typically overcomes the thermal stability of certain sorbents, such as metal halides, which decompose above about 400°C. Therefore, in some other preferred embodiments, it may be desirable to pre-reduce the catalyst in the reactor before integration with the sorbent, thereby eliminating the need for a reduction procedure in the presence of an sorbent that decomposes at the reduction temperature.

[0118] In some other embodiments, the decomposition of the adsorbent may be amplified by trace amounts of water molecules, particularly those that spontaneously absorb to the metal halide in the presence of moisture, forming metal halide hydrates, e.g., MnCl2-xH2O. Because NH3 binds more strongly to the adsorbent than water, the adsorbent may be subjected to ammonia before the active catalyst is supported or mixed, resulting in the ammonia being absorbed by the adsorbent before any high-temperature (above 300°C) process. By subjecting the adsorbent to ammonia, the ammonia is absorbed, replacing previously absorbed water, thus increasing the stability of the adsorbent. The form of the catalyst-adsorbent structure may include pellets, tablets, extruded products, or granules, and is generally understood to encompass any structure that flows freely while supported in a reactor.

[0119] Catalyst-sorbent structures in the form of pellets, tablets, extruded products, or granules may have an average diameter of 1 mm to 20 mm, preferably 3 mm to 10 mm, and more preferably 3 mm to 9 mm.

[0120] A catalyst-adsorbent structure having an outer coating shell may have an outer coating shell with an average thickness of about 3 microns to about 200 microns, preferably about 10 microns to about 150 microns, and more preferably about 20 microns to about 100 microns.

[0121] A catalyst-adsorbent structure having one or more inner coating shells may have inner coating shells with an average thickness between about 3 microns and about 200 microns, preferably between about 10 microns and about 150 microns, and more preferably between about 20 microns and about 100 microns.

[0122] A catalyst-adsorbent structure in the form of a monolithic structure may include a support material. The support material may preferably be a ceramic material, a metal oxide such as alumina, or a combination thereof. The support material may be microporous, macroporous, or a combination thereof. The support monolith may be impregnated, coated, or otherwise supported on the active catalyst portion and / or the adsorbent portion. The monolithic structure may be partially or completely formed from an adsorbent material such as a zeolite, a metal halide, or a combination thereof. In exemplary embodiments in which the monolithic structure is formed from an adsorbent material, the monolithic structure may be impregnated, coated, or otherwise supported on the active catalyst portion and / or the adsorbent portion, thereby being an additional adsorbent material which may be the same adsorbent as the monolithic structure or a different adsorbent than the monolithic structure.

[0123] In some embodiments, monolithic structures formed from metal oxides or zeolites can be modified by ion exchange to incorporate metal cations, alter surface properties, and adjust affinity for ammonia adsorption. Furthermore, monolithic structures can be silanized, hydrated, doped, or modified in similar manner to alter the hydrophobicity of the structure. The monolithic structure may be of a metal variety and may be coated, impregnated, fused, or otherwise supported on the active catalyst portion and / or sorbent portion. Metal monolithic structures may have higher thermal conductivity than ceramic, metal oxide, or zeolite monolithic structures. Those skilled in the art will understand that higher thermal conductivity can be advantageous by rapidly heating or cooling the monolithic structure, and therefore the active catalyst portion and / or sorbent portion. It can be further understood that monolithic structures can simplify the reactor support process by supporting a single structure rather than multiple structures. Monolithic structures can also reduce pressure loss across the entire activated bed by increasing the total void ratio and decreasing flexibility.

[0124] In some embodiments, the active catalyst and / or sorbent portion supported on the monolithic structure may take the form of metal ions, or nanocrystals, or microcrystals. Any such monolithic structure may be highly ordered, particularly by zeolites, by extrusion, addition, or other manufacturing methods that induce structured 3D periodicity throughout the structure. In some exemplary embodiments, the structured 3D periodicity throughout the structure may have the form of a honeycomb containing hundreds to thousands of parallel channels or pores defined by many thin walls within the honeycomb structure. The channels may be square, hexagonal, circular, or other shapes. The pore density is 1 cm². 2 The density may be between 30 and 200, and the separation wall may be approximately 0.05 mm to 3 mm thick.

[0125] The catalyst-sorbent structure preferably has a partial sorbent load of 5% to 95% by weight, preferably 10% to 90% by weight, and more preferably 20% to 90% by weight. The catalyst-sorbent structure preferably has a partial active catalyst load of 0.01% to 20% by weight, preferably 0.25% to 10% by weight, and more preferably 0.5% to 5% by weight. In some preferred embodiments, the catalyst-sorbent particles have a partial active catalyst load of less than 5% by weight. The catalyst-sorbent structure preferably has a (catalyst:sorbent) weight ratio of catalyst load to sorbent load of about 1:1 to about 1:300, preferably about 1:1 to about 1:50, and more preferably about 1:1 to about 1:10.

[0126] The predicted partial loading density of the sorbent is approximately 100 kg / m³. 3 ~2000kg / m 3 The range is preferably about 300 kg / m 3 ~Approx. 1500kg / m 3 Within the range of approximately 500 kg / m², more comfortably. 3 ~Approx. 1200kg / m 3 It can be within this range. The predicted catalyst partial loading density is approximately 100 kg / m³. 3 ~2000kg / m 3 The range is preferably about 300 kg / m 3 ~Approx. 1500kg / m 3 Within the range of approximately 500 kg / m², more comfortably. 3 ~Approx. 1200kg / m 3 It is within the range of [the specified range].

[0127] In some embodiments where an increased cycle time is desired, the weight load of the catalyst portion is approximately 5 kg / m³. 3 ~about 500kg / m 3 Preferably about 5 kg / m 3 ~about 400kg / m 3 Comfortably about 5 kg / m 3 ~about 250kg / m 3 While it can be within this range, the weight load of the adsorbent portion is approximately 50 kg / m 3 ~1500kg / m 3 Preferably about 150 kg / m3 ~Approx. 1400kg / m 3 More preferably about 250 kg / m 3 ~Approx. 1200kg / m 3 The range can be as follows. In some embodiments, the (catalyst:adsorbent) weight ratio of the catalyst portion to the adsorbent portion can be in the range of approximately 1:3 to approximately 1:300.

[0128] Processes involving the integrated catalyst-sorbent structure of this disclosure exceed the equilibrium NH3 composition obtained without the use of the sorbent portion. In some embodiments, the nitrogen conversion rate (in terms of the percentage of stoichiometric conversion of unreacted nitrogen feedstock to ammonia) may be in the range of about 30 to 99.99%, preferably 50 to 99.9%, and more preferably about 70 to 99%, per pass, where a pass is understood to be a single tube or a series of segmented tubes that are fluidly connected during one process feed cycle. In some embodiments, the hydrogen conversion rate (in terms of the percentage of stoichiometric conversion of unreacted hydrogen feedstock to ammonia) may be greater than 70% per pass, in some embodiments at least 80% to a maximum of 100%, in some other embodiments at least 80% to a maximum of 99.99%, and in some other embodiments at least 80% to a maximum of 99%, where a pass is understood to be a single tube or a series of segmented tubes that are fluidly connected during one process feed cycle.

[0129] The integrated catalyst-sorbent structure of the present disclosure comprises a catalyst portion and an sorbent portion, wherein the catalyst portion can convert unreacted hydrogen feedstock and unreacted nitrogen feedstock into ammonia products, and the sorbent portion can absorb the generated ammonia; both the conversion of the catalyst portion and the absorption of the sorbent portion can occur at temperatures in the range of about 100°C to about 500°C, preferably about 200°C to about 400°C, more preferably about 250°C to about 350°C, and even more preferably about 280°C to about 330°C; both the conversion of the catalyst portion and the absorption of the sorbent portion can occur at pressures in the range of about 2 bar to about 200 bar, preferably about 5 bar to about 100 bar, more preferably about 5 bar to about 50 bar, and even more preferably about 5 bar to about 20 bar.

[0130] Ammonia production and capture using catalyst-sorbent structures The disclosure also relates to a process for producing ammonia, the process comprising providing the catalyst-sorbent structure of the disclosure as a fixed bed, preferably a packed bed, in a reactor, under normal operating conditions, the catalyst portion converts unreacted hydrogen and unreacted nitrogen feedstocks into ammonia products, and the sorbent portion captures the generated ammonia.

[0131] The catalyst-sorbent structure is preferably placed in the reactor and is supported in a range of about 0.1% to about 99.9% of the reactor volume, preferably about 10% to about 80%, preferably about 15% to about 60%, more preferably about 20% to about 40%, and in some embodiments even more preferably about 25% to about 35%. When packed, the catalyst-sorbent pellets of the present invention are generally understood to have an interstitial void volume in the range of about 20% to 50%. The defined loading amount in the reactor volume defines the portion of the reactor volume containing the packing material without considering the void volume.

[0132] The amount of catalyst-sorbent supported in the reactor is preferably at least 10% of the reactor volume, in some embodiments at least 20%, in some embodiments at least 30%, in some embodiments at least 40%, in some embodiments at least 50%, in some embodiments at least 60%, in some embodiments less than 95%, in some embodiments less than 80%, and in some embodiments less than 70%.

[0133] The catalytic portion of the catalyst-sorbent structure can be present in the reactor in a weight range (w / w) of about 0.01% to about 20%, preferably about 0.25% to about 10%, and more preferably about 0.5% to less than 5%. In some preferred embodiments, the catalyst-sorbent particles have a catalyst portion load in the reactor of less than 5% by weight.

[0134] The sorbent portion of the catalyst-sorbent structure can be present in the reactor in a weight range (w / w) of about 5% to about 95%, preferably about 10% to about 90%, and more preferably about 20% to about 80%.

[0135] The catalyst-sorbent structure can be present in the reactor in a catalyst-sorbent weight ratio of about 1:1 to about 1:300, preferably about 1:10 to about 1:50, and more preferably about 1:15 to about 1:25.

[0136] In some preferred embodiments, the process for producing ammonia using the catalyst-sorbent structure of the present disclosure has a process cycle that is less than the full sorbent capacity of the sorbent portion. In some embodiments, the process cycle is at least 20% to about 95% of the full theoretical capacity, as defined by the operating temperature and operating pressure of the process bed.

[0137] A process for producing ammonia using a catalyst-sorbent structure may have an initial process cycle having an initial transformation and a second process cycle having a second transformation, the second transformation having a lower transformation than the initial transformation, in some embodiments at least 0.1% lower than the initial transformation, and in some preferred embodiments 1% to 10% lower than the initial transformation.

[0138] A process for producing ammonia may include providing the catalyst-adsorbent of the present disclosure to multiple beds. In some embodiments, the multiple beds are provided in series. In some embodiments, the multiple beds are provided in parallel. In some embodiments, the multiple beds are provided in both series and parallel configurations.

[0139] Unreacted hydrogen can be supplied from a hydrogen source. While it is intended that unreacted hydrogen can be supplied from any hydrogen source, in some preferred embodiments the hydrogen source includes production from water in an electrolytic cell.

[0140] Unreacted nitrogen can be supplied from a nitrogen source. While it is intended that unreacted nitrogen can be supplied from any nitrogen source, in some preferred embodiments the nitrogen source is a pressure swing adsorption (PSA) system, an air separation unit (ASU) system, a membrane separator, or a combination thereof.

[0141] Generally, the process of this disclosure is operated in a circulating mode, thereby flowing unreacted feedstock of hydrogen and nitrogen over a catalyst-sorbent structure in a first bed, thereby causing the catalyst portion to produce ammonia, and thereby the sorbent portion to capture the produced ammonia. The sorbent may include an absorbent, an adsorbent, or a combination thereof. Before the ammonia is discharged or a breakthrough is achieved at the end of a tube containing the closely contacting sorbent and catalyst, the process feed is switched to a second bed containing the closely contacting catalyst and sorbent, thereby allowing the process to continue forming ammonia. The ammonia may fill only a portion of the first integrated reactor-absorber section at a capacity of approximately 10–99% of its theoretical maximum at a given temperature and pressure. After the feed is switched to a single or series-connected partitioned tube of the regenerated integrated reactor and absorber, the ammonia absorbed in the previous cycle is desorbed. The desorption process is carried out to remove ammonia from the system and minimize the reverse reaction, resulting in more than 80% of the ammonia generated in the sorbent portion of the process being captured during desorption, preferably more than 90%, and more preferably more than 95% of the generated ammonia being recovered.

[0142] Generally, absorbents react with ammonia according to formula (2) as follows: [Absorbent] + NH3 → [Absorbent] - xNH3(2) Here, several values ​​for "x" are possible depending on the chemical properties of the sorbent material, temperature, and pressure. The sorbent can be supported on a porous material. The sorbent can absorb ammonia mainly into the bulk of the material, first by absorption into surface molecules. The sorbent can also capture and store ammonia mainly at the interface between the sorbent and the gas phase through chemical or physical sorbent or bonding.

[0143] After desorption, the gas stream can contain ammonia of substantially high purity. In some embodiments, the gas stream can contain ammonia in nitrogen gas ranging from about 0 vol% NH3 to about 100 vol% NH3. In some other embodiments, the gas stream can contain ammonia in a mixture of H2 and N2, where the ammonia is in the range of about 5 vol% to 99 vol%, preferably about 10% to 99%, more preferably about 20% to 99%, more preferably about 30% to 99%, more preferably about 40% to 99%, more preferably about 50% to 99%, and most preferably the ammonia is greater than about 50%.

[0144] The gas stream can be pressurized and / or cooled according to the vapor-liquid equilibrium of the mixture and condensed into substantially pure NH3 as defined by a mass fraction of over 90%, preferably about 90% to about 100%, more preferably about 92% to about 100%, more preferably about 94% to about 100%, more preferably about 96% to about 100%, and even more preferably about 98% to about 100%. The remaining gas can be N2 with trace amounts of ammonia, as determined from the vapor / liquid equilibrium. It is beneficial to select condensation conditions such that the gas stream exiting the condenser unit has minimal NH3. This mostly N2 gas stream can be used, for example, through a boost condenser to repressurize the process after a lower-pressure desorption step and reused as a sweep gas for subsequent or parallel desorption steps.

[0145] Alternatively, pressure swing adsorbents can be used downstream of a reactor or combined reactor and adsorbent to purify ammonia.

[0146] In this disclosure, it should be understood that desorption occurs from an sorbent (e.g., an absorbent and / or adsorbent) by increasing the temperature of the sorbent and / or decreasing the partial pressure of NH3 in the sorbent and / or decreasing the total pressure in the sorbent. It is beneficial to operate sorbent / desorption substantially isothermally, which is defined as a thermal gradient of the bed from about 0°C to about 10°C, so that heat can be transferred between modes, otherwise desorption may require higher temperatures, and heat exchange using a liquid heat transfer fluid becomes more complex. In some embodiments, a substantially isothermally isothermally pressure swing is desired, so that the bed remains isothermally or nearly isothermally in each cycle, such as within about 30°C, more preferably within about 20°C, and even more preferably within about 10°C, and the pressure decreases from adsorption mode to desorption mode, causing the elution of previously adsorbed NH3.

[0147] To simplify the downstream NH3 condensation / separation step, it may be beneficial to maximize the desorption pressure. It is beneficial that the desorption pressure is higher than the NH3 condensation pressure at or near ambient temperature, thereby eliminating the need for further pressurization for ammonia condensation at ambient temperature. In some preferred embodiments, the desorption pressure is at least about 1 bara, in some embodiments at least about 2 bara, in some embodiments at least about 5 bara, in some embodiments at least about 8 bara, in some embodiments at least about 11 bara, in some embodiments at least about 13 bara, and in some embodiments at least about 15 bara higher than the NH3 condensation pressure at ambient temperature, and the NH3 condensation pressure is preferably in the range of about 4 bara to about 15 bara at about -15°C to about 30°C. It may be beneficial to have the highest possible pressure for ammonia desorption, thereby minimizing the cooling (or chilling) required to enable NH3 condensation under desorption pressure, or minimizing recompression to condense NH3 at a given temperature.

[0148] Maximizing the desorption pressure to minimize the reverse reaction of the generated ammonia to N2 and H2 may be beneficial, and this is thermodynamically more preferable at low pressures.

[0149] Minimizing the desorption pressure can be beneficial, as it allows a larger working capacity of the sorbent to be available in the circulation process. In some embodiments, the desorption gas stream pressure is below the pressure of NH3 condensation at a given temperature, and as a result, the desorption stream containing NH3 can be compressed to favor NH3 condensation in the subsequent cooling and / or condensation steps.

[0150] In some embodiments following the pressure swing, the desorption bed can be at a pressure significantly lower than the pressure of the supply gas supplied to the parallel bed operating in sorption mode, which in some embodiments is preferably adsorption mode. The total pressure is preferably 3 bar or more below the supply pressure, in some embodiments 5 bar or more below the supply pressure, in some embodiments 7 bar or more below the supply pressure, in some embodiments 9 bar or more above the supply pressure, in some embodiments 11 bar or more above the supply pressure, in some embodiments 13 bar or more above the supply pressure, and in some embodiments it can be about 15 bar below the supply pressure. In some embodiments, the total pressure is about 1 to about 20 bar lower than the supply pressure, and in some other embodiments, it is about 5 to about 15 bar lower than the supply pressure. Unreacted effluent from the bed in sorption mode can be supplied to the bed that has just undergone desorption in order to repressurize the bed. Similarly, during the switch from adsorption mode to desorption mode, there may be an initial depressurization event in which unreacted feed gas is rapidly discharged from the floor head space and gaps. This unreacted feed gas can be reused from N2 or unreacted gas discharged during the exhaust step, but preferably does not contain NH3. These unreacted feed gases can also be supplied to the recently depressurized floor to repressurize it. These unreacted feed gases can also be used as scavenging gases during the desorption step on a parallel floor. Scavenging gases containing H2 can suppress the reverse reaction / decomposition of desorbed ammonia by reducing the equilibrium propulsion force of the reverse reaction.

[0151] In some embodiments relating to the desorption rate in zeolites, the desorption rate may be limited by the mass transfer reaction rate rather than the bulk desorption reaction rate, as in metal halides, thereby the rate limiting step is typically the diffusion of NH3 (e.g., as ions) from the crystalline structure to the surface. Alternatively, in the desorption rate of zeolites, the mass transfer reaction rate relates to desorption (e.g., as gas molecules) from the inside to the outside of the pores. In the process of the present invention, where the adsorbent is a zeolite or a combination of zeolites, desorption may be the fastest step in the process cycle. Thirdly, a repressurization step may be beneficial. In this step, H2 and N2 can be supplied to a recently depressurized bed, and the bed can be filled with unreacted and / or inert gases to restore the pressure to the desired reaction pressure.

[0152] In some embodiments, it may be advantageous to maximize the desorption rate to minimize the residence time of desorbed NH3 in the reactor and prevent the reverse reaction. The decomposition reaction catalytically decomposes the product NH3 back into N2 and H2, and higher temperatures and higher catalyst loads are preferred. Reducing the reaction time or residence time spent in contact with the catalyst by the desorbed ammonia is essential to maximize NH3 recovery. As shown in Table 13, a surface velocity exceeding approximately 0.01 m / sec and a catalyst load density of approximately 100 kg / m³ are desirable. 3 ~about 500kg / m 3 When using an iron-based catalyst at 3 bara, the predicted decomposition conversion rate of NH3 is less than approximately 1.11% at temperatures below approximately 330°C, and the surface velocity is defined by dividing the volumetric flow rate at the actual temperature and pressure by the cross-sectional area of ​​the open channel. For alternative catalyst compositions, the critical surface velocity, temperature, and pressure may differ, but are determined based on the reaction rate, catalyst loading, and decomposition conditions in order to maintain a performance of less than 10% decomposition of NH3 during the decomposition step.

[0153] In certain embodiments, the process of the present invention includes an additional fluid chamber positioned within the packed bed containing the reaction and sorbing, such that the fluid flow is substantially radial rather than axial during desorption. The reactor configuration of the present invention is first manufactured comprising an inner porous tube, an annular gap in which the sorbent and catalyst of the present invention can be contained therein, and an outer heat transfer wall for removing the exothermic reaction and adsorption / absorption. After manufacturing, the catalyst-sorbent structure, such as in pelletized form, is supported in the annular flow region in a manner consistent with conventional best practices for catalyst support in tubular or annular reactor chambers. As illustrated in Figure 11, the porous tube or chamber is maintained with a closed outlet during reaction and sorbing and an open outlet during desorption, so that desorbable NH3 spends the shortest possible time near the catalyst, in order to minimize the decomposition reaction. Otherwise, the length of the reactor bed containing the catalyst, which the desorbed ammonia must traverse, is reduced by an order of magnitude or more for the radial flow desorption step, thereby allowing the product ammonia to be removed more rapidly from the catalyst and inhibit the decomposition reaction. The radial flow length may be in the range of approximately 0.02 m to approximately 0.1 m, representing a packed cyclic catalyst-sorbent flow chamber. The reactor length may be in the range of approximately 1 m to approximately 15 m, such that the ratio of the axial flow length to the radial flow length of the reactor is approximately 10 to approximately 750. During the sorption step, the reactive species flow along the axial flow length, while during the desorption step, the product species flow along the radial flow length. The decomposition time is reduced approximately proportionally to the reduction in the flow length of the product species through the desorption process when in contact with the catalyst-sorbent particles of the present invention.

[0154] Catalyst-sorbent structure and process configuration Integrated catalyst and sorbent systems can offer better performance than recycling reactors, and comparative recycling reactors supported with ammonia synthesis catalysts at or near theoretical packed bed densities are operated in recycling mode. Furthermore, recycling reactors require substantial recompression and reheating of the unreacted feed mixture at the reactor inlet.

[0155] In some embodiments, it may be desirable for the weight fraction of catalyst per reactor volume to be lower than that of the sorbent. A lower weight load is required for a sufficiently active catalyst, while a higher weight load of sorbent increases the feed switching cycle time while the feed is moved to a fresh reactor tube to continue the process. After switching the feed to the second bed, the absorbed ammonia is then desorbed on the first bed. A lower catalyst load can reduce the rate of the reverse reaction during desorption. In some embodiments, the weight load of the active catalyst is approximately 5 kg / m³. 3 ~500kg / m 3 It may be within this range, and the weight load of the adsorbent should be approximately 50-1500 kg / m³. 3 The range may also be within this range. The ratio of the sorbent to the catalyst may be in the range of about 3 to about 300.

[0156] In some embodiments, the nitrogen conversion rate per pass (in terms of stoichiometric conversion percentage) for forming ammonia in the presence of hydrogen is at least about 30%, in some embodiments about 30% to about 99%, and in some other embodiments about 50% to about 99%. In some embodiments, the hydrogen utilization rate in combination with the catalyst-sorbent structure is over 70%, and in some embodiments about 80% to about 99%.

[0157] In relation to the ratio of cycle time to reactor volume, it may be desirable to decrease the process volume while increasing the cycle time. Reactor volume can be minimized by having more catalyst and less sorbent, which allows for faster filling of the reactor volume for desorption and reuse. With less catalyst and more sorbent, the reactor volume fills more slowly, allowing for longer cycle times. Longer cycle times are desirable due to longer catalyst and sorbent life, less mechanical wear of valves, and reduced feed loss remaining in the open process volume observed in the catalyst-sorbent structural gap, between catalyst-sorbent structures, and in the header and footer during each cycle at feed switching.

[0158] In relation to cycle times below full sorbency capacity, the process may be circulated when the absorbent is at least about 20%–95% of its full theoretical capacity, as defined by the operating temperature and pressure of the process bed. As the bed begins to fill forward towards the reactor absorber, the conversion rate decreases. The process may run at an initial conversion rate of about 90%–99%, which decreases to at least a second conversion rate that is at least about 1% or 1%–10% lower than the first conversion rate. In some embodiments, the stability of the absorbent may be improved by circulating it between limited capacity ranges and for higher coordination numbers defined by the transfer of 2–1 or 6–2 moles of ammonia per mole of absorbent.

[0159] In relation to the operating temperature and pressure, it may be desirable to increase the temperature for a faster reaction rate, as defined by the moles of ammonia per mole of absorbent or the grams of ammonia retained per gram of absorbent, while simultaneously minimizing the temperature to obtain a higher adsorption capacity. It may also be desirable to operate at a temperature lower than the phase transition temperature of the absorbent. In one embodiment, the temperature is set to allow for high coordination number absorption (e.g., 6 to 2 moles of NH3 per mole of absorbent transition in MnCl2, which is four times the capacity and half the heat generation of the 2-1 transition). The reactor temperature is defined by the wall temperature or maximum temperature in at least a portion of the reactor and absorbent bed of the present invention. In some embodiments, higher pressure NH3 also allows for a higher absorption capacity. The selection of reaction and desorption pressures to match the upstream feed compression needs and downstream ammonia compression needs must minimize process volume, metal weight, heat input energy, and compression energy. In some embodiments, pressure and temperature swings must be considered. Isothermal or nearly isothermal with pure pressure swings allows for the reuse of virtually all heat. Tubes or other heat transfer regions may be connected between beds in opposite modes so that sorption and desorption can occur at the same temperature and the heat of desorption provided by the heat of sorption is released. Small swings are preferred in terms of time and energy. Larger swings may be preferred considering the trade-offs between reaction rate and capacity. In some embodiments, consideration of scavenging gases can minimize dilution of NH3 during desorption while minimizing reverse decomposition or ammonia decomposition reactions. In some embodiments, an electric heating element in at least a portion of the tube length can improve the rate of desorption or utilize an intermittent energy source. In some embodiments, the catalyst-sorbent structure may be heated directly. In other embodiments, the catalyst-sorbent bed may be heated or cooled through the tube wall.

[0160] With catalyst-sorbent structures having low catalyst loads, in relation to the temperature and pressure profiles, the pressure decreases as the number of moles of ammonia sorbent produced is sorbed in or on the solid absorbent, and the heat transfer of the packed bed degrades with low-density gases, increasing the thermal gradient. There may be a balance between catalyst load, rate, operating pressure, and temperature that creates a combination where the reactor is thermally controlled or has high-temperature hot spots that sinter the catalyst or absorbent or reduce the absorbent's ability to retain ammonia.

[0161] In some embodiments, there exists a segmented process in which two or more reactor absorbent beds can be operated in series so that each section is independently optimized by adjusting parameters such as length, diameter, catalyst load, rate, temperature, and particle diameter. In some embodiments, the segmented process may be operated so that the desorption of a second or subsequent section occurs after the desorption of the first section. In some embodiments, the desorption of a second or subsequent section may occur (begin or end) after feed switching or a cycle to the first segment.

[0162] In some embodiments, the catalyst-sorbent structure has a redox activation protocol. In some embodiments, one or more redox activation steps are performed before operation. In some embodiments, redox activation can be performed in both steps of a stepwise reactor, or in only the first or subsequent steps of a series reactor, to increase the activity over a portion of the reactor length. For example, cobalt is reduced at about 350°C to about 450°C, preferably in the range of about 395°C to about 420°C. In some embodiments, the oxidation step of the redox can be performed in excituated before loading and operation, so that only the final reduction step in diluted hydrogen is completed before carrying out the ammonia synthesis reaction.

[0163] In some embodiments, there is a designed interaction between the catalytic and sorbent portions. For example, the sorbent portion may function as a catalyst promoter, and if strongly interacting, may provide overflow sites for intermediates from the catalyst surface, potentially influencing the apparent reaction order for H2, N2, or NH3. Similarly, NH3 generated on the catalyst surface may be preferentially adsorbed by the sorbent portion, potentially through a ripple mechanism, altering the apparent reaction order of NH3. In one embodiment, the apparent reaction order is regulated by adsorbed species with higher surface concentrations competing with the catalytic site for conversion. This effect may decrease the net reaction rate by blocking sites or increase the net reaction rate by eliminating the blocking site species. This same effect may increase the sorption rate of NH3 to the absorbent. The presence of one component may increase the site density of the other, increasing the surface area-to-volume ratio and accelerating the reaction rate. In general, the affinity of an sorbent for NH3 can alter the reaction sequence of the catalyst toward NH3, and co-supported absorbents / catalysts may perform better than independently supported mixtures due to molecular interactions. In some embodiments, it may be beneficial in this disclosure to increase the apparent reaction rate that produces NH3 so that the reverse reaction is minimized by limiting the adsorption of NH3 or by the presence of other species that remove NH3 after it has formed on the catalyst site. The forward reaction is maximized by the tendency of NH3 to move from the catalyst surface to nearby absorbent sites, thereby generating a molecular mass transfer propulsion force. In some embodiments, the interaction between the co-supported catalyst and the absorbent means an optimal loading amount that maximizes performance, for example, performance may increase with increasing catalyst loading (assuming a constant absorbent loading amount) until it decreases or plateaus.

[0164] In some embodiments, the sorbent portion may have an activation protocol. The sorbent may be partially activated after the catalyst-sorbent structure is supported on the reactor by subjecting the catalyst-sorbent structure to a temperature above the temperature of the sorbent hydration reaction (i.e., MY-xH2O → MY+H2O) (40°C to over 400°C, depending on the absorbent) in a substantially dry process gas free of water vapor, as defined by a vapor pressure of less than about 10 millibars. In some embodiments, the sorbent may be subjected to a formation cycle, where the ammonia sorbent / desorbent process occurs at least once through at least a desired volume range, sometimes beyond the desired volume range, with grinding or other mechanical changes from the cycling occurring before final integration in the combined system, and the form of the absorbent is set before use in the combined system. In some embodiments, the sorbent remains unaffected by the reduction or redox activation steps necessary to generate a substantially reduced metal catalyst portion for ammonia synthesis in close or in close contact with the absorbent material.

[0165] Multiple floors in series, parallel, or series and parallel may be used to enable operation at low production capacity obtained from variable renewable energy sources. It may be desirable to connect the process to an energy source that generates time-varying or intermittent amounts of electricity, as seen from renewable sources such as wind, solar, tidal, or geothermal. Hydrogen and / or nitrogen production or purification sources may be operated by such intermittent power. In some embodiments, the hydrogen source is an electrolytic cell. In some embodiments, the nitrogen source is a PSA system, or ASU system, or membrane separator, which may be electrically driven. If it is desirable to operate with less power than the designed capacity, the nitrogen and / or hydrogen sources may be operated at partial capacity or reduced operation below plant nameplate capacity so that the available feedstock is just enough to produce a certain amount of the designed NH3 capacity. In some embodiments, one or more of the nitrogen or hydrogen sources may be operated at full capacity, while the other is operated below full capacity, and any excess product may be stored or released.

[0166] In some embodiments, if it is desirable and / or necessary to generate less than the designed NH3 capacity, the feed material flow to one or more floors may be stopped so that the floors are maintained in a high-temperature, warm-temperature, or low-temperature standby mode so that idle units are not in operation. Low-temperature standby is defined by lowering the unit process temperature to ambient conditions. Warm-temperature standby is defined by an intermediate temperature between the ambient temperature and the overall process operating temperature. High-temperature standby is defined by maintaining a target operating temperature while not in use, and the process can be restarted quickly, defined by restarting for a time ranging from a few seconds to a few hours, preferably in the range of 10 seconds to 60 minutes.

[0167] In some embodiments, the tubing containing the floor can be exhausted to remove process gases by reducing the internal pressure through decompression or by a vacuum pump. In some embodiments, the floor can be left isolated during standby mode. The remaining substantially stagnant gas in the floor may react to form NH3, which may be absorbed by an absorbent. In some embodiments, the heat generated can be actively removed or passively lost due to the system's heat loss. In other embodiments, the floor can be thermally isolated when the feed flow is stopped to reduce the cooling load. The heat of the system can create a sink for collecting heat generated by the substantially stagnant feed present in the system during standby mode.

[0168] In some embodiments, the dormant reactor sorbent bed may remain in a variety of absorbent states, including substantially cleaned, defined as less than 10% of the theoretical capacity; partially absorbed, defined as about 10% to about 80% of the theoretical capacity; or completely absorbed, defined as about 80% to 100% of the theoretical capacity. The dormant process may be restarted after a dormancy period ranging from less than 1 second to a maximum of 15 days, preferably about 10 seconds to about 16 hours. Restarting may occur by either the reaction and sorbent parts of the process cycle, or through desorption to remove adsorbed ammonia remaining in the hardware. Restarting the operation through combined reaction and sorbation may improve the start-up time by leveraging the exothermic reaction to quickly heat to the desired temperature. In one embodiment, the combined flow to the reactor and absorber may differ from steady-state cycle operation, as there is more exothermic heat generated at certain locations, raising the unit operation temperature from a low-temperature or warm-temperature standby mode.

[0169] In some embodiments, the system shall consist of valves and piping such that the floors can be connected or disconnected in series or in parallel without modification of the apparatus, so that changes can be made on a short time scale, defined as about 10 seconds to 60 minutes, preferably less than about 5 minutes.

[0170] The process of the present invention closely integrates the catalyst formation of ammonia from nitrogen and hydrogen, accompanied by the separation of ammonia, thereby reducing the thermodynamic limits of the conversion. Close contact between the catalyst and the absorbent is achieved by achieving proximity of molecules of both materials so that they are retained within a single pellet or other structural component, as disclosed herein.

[0171] In some embodiments, at least one reactor bed contains a catalyst-sorbent structure. Preferably, the reactor contains two or more reactor beds containing catalyst-sorbent structures, and more preferably, the reactor contains multiple reactor beds containing catalyst-sorbent structures. In some embodiments, the composition of the catalyst-sorbent structure in each reactor bed is substantially the same with respect to the axial length and width of the reactor bed. In some other embodiments, the composition of the catalyst-sorbent structure varies below the length of the bed in at least one reactor bed, in some embodiments two or more reactor beds, and in some other embodiments each of the reactor beds.

[0172] In an exemplary embodiment, the amount of catalyst supported may be higher at the front of the reactor bed than at the rear of the reactor bed. In another exemplary embodiment, the amount of catalyst supported may be higher at the rear of the reactor bed than at the front of the reactor bed. In yet another exemplary embodiment, the amount of catalyst supported may be higher in the middle of the reactor bed than at the front or rear of the reactor bed.

[0173] In an exemplary embodiment, the amount of sorbent may be higher at the front of the reactor bed than at the rear of the reactor bed. In another exemplary embodiment, the amount of sorbent may be higher at the rear of the reactor bed than at the front of the reactor bed. In yet another exemplary embodiment, the amount of sorbent may be higher in the middle of the reactor bed than at the front or rear of the reactor bed.

[0174] In some embodiments, multiple reactor beds may be implemented in parallel within the reactor, and the composition of the catalyst-sorbent structure in each reactor bed may differ. In some embodiments, in a multiple reactor bed configuration, the composition of the catalyst-sorbent structure in one or more reactor beds may vary along the axial length, width, or a combination thereof.

[0175] In further embodiments, a particular reactor bed may have different relative amounts of catalyst:sorbent. In exemplary embodiments, one or more reactor beds may have 100% catalyst, while one or more other reactor beds may have 100% sorbent. In such beds, multiple catalysts or sorbents may be used and may vary along the length of the bed. In some preferred embodiments, reactor beds having only catalyst or primarily catalyst are configured in the reactor to be located before the other reactor beds in relation to the introduction of unreacted nitrogen and unreacted hydrogen feedstocks. In some preferred embodiments, reactor beds having only sorbent or primarily sorbent are configured in the reactor to be located after the other reactor beds in relation to the introduction of unreacted nitrogen and unreacted hydrogen feedstocks.

[0176] In exemplary embodiments, a first reactor bed is located before other reactor beds, having a catalyst-only loading amount greater than that of other reactor beds, in relation to the introduction of unreacted nitrogen and unreacted hydrogen feedstocks; a third reactor bed is located after other reactor beds, having an sorbent-only loading amount greater than that of other reactor beds, in relation to the introduction of unreacted nitrogen and unreacted hydrogen feedstocks; and a second reactor bed is located between the first and third reactor beds, having a mixed loading amount of catalyst and sorbent, for example, the catalyst-sorbent structure of the Disclosure. In some embodiments, the second reactor bed comprises one or more reactor beds having the catalyst-sorbent structure of the Disclosure.

[0177] This disclosure envisions a process of the present invention in which a first cycle, including sorption and reaction, is thermally connected to a desorption process. In one embodiment, concentric tubes may be used, which are filled with a catalyst-sorbent structure that maintains close contact. In the first part of the cycle, feed may enter a first bed in which exothermic reaction and sorption occur. Heat generated from the first process is transferred to an adjacent chamber in which desorption occurs. The process may be configured as a concentric reactor bed such that one bed is maintained within a central core and a second bed is maintained within an annular region surrounding the first bed. In an alternative embodiment, a heat exchange chamber may be located in the core such that both the first and second processes are configured as an annular packed bed. In this embodiment, the reactor volume and total amount of metal are reduced by sharing walls between the reaction cycle and the desorption cycle. Furthermore, the proximity of the two processes allows for minimized reactor header and footer volumes to enable higher utilization of hydrogen and nitrogen reaction feed materials.

[0178] The desorption energy may be further enhanced by the use of resistance heating, thereby the heating element may be embedded in the wall at one or more axial positions to further modify the temperature so that the absorption coordination number decreases and the ammonia absorbed in the previous cycle is desorbed more rapidly and easily.

[0179] This disclosure envisions a system of the present invention for cobalt-containing pellets or related structured catalysts that can achieve increased activity by using two or more redox activation steps prior to operation.

[0180] Therefore, the catalyst-sorbent structures of this disclosure may undergo redox activation before being used for combined ammonia synthesis and sorbation. The catalyst-sorbent structures may undergo sequential redox oxidation before being supported and operated in the plant, such that only the final reduction step must be performed in situ before operating the plant to produce ammonia. The total number of redox steps may be two or more for at least a portion of the plant. In one embodiment, a portion of the catalyst-sorbent structure supported in one or more steps is redox activated in two or more steps, while a portion of the catalyst-sorbent structure in one or more steps has one or fewer redox activation steps. The effective catalytic activity per gram of catalyst in one or more steps in this manner is higher than the effective catalytic activity per gram of catalyst in one or more steps.

[0181] In one embodiment, the catalyst is more active in the first step than in the second step. In an alternative embodiment, the catalyst is less active in the first step than in the second step.

[0182] It is desirable to avoid or minimize the catalyst oxidation step within the plant, and it is desirable to support a pre-oxidized catalyst for final in-situ reduction. The catalyst of the present invention may also be reduced in situ before operation or periodically during operation, depending on the need for catalyst regeneration with a hydrogen-containing feed gas.

[0183] This disclosure intends that combined catalysts and absorbents can be deployed in a desired configuration. Unreacted nitrogen and unreacted hydrogen in the first part of the process can be used to improve the desorption of ammonia in the second part of the process cycle by acting as scavenging gases to reduce the desorption time. The flow rate of the unreacted feed can be combined from two or more pipes to increase the flow rate to a single pipe, which increases the desorption rate or reduces the average residence time for desorption.

[0184] The combined unreacted feed from the first part of the cycle can sequentially desorb ammonia from the parallel tubes. As shown in Example 5 below, the higher rate during desorption reduces the residence time, which reduces the amount of ammonia lost to decomposition. In the desorption process of the present invention, multiple tubes configured in parallel that were previously operated for combined reaction and sorbation in one cycle can be desorbed in series or series-parallel configuration. In the example having 100 tubes operated in parallel between the reaction and sorbent parts of the process, then, for example, the first 10 tubes are desorbed, followed by the second 10 tubes, and so on, providing a higher rate during desorption than during reaction and sorbation.

[0185] In one embodiment, the inlet surface velocity during desorption is greater than 1.2 times the outlet surface velocity during the combined sorption and reaction, such that the desorption tube is processed in a partially sequential mode. In another embodiment, the inlet surface velocity during desorption is 1.2 to 1500 times higher, preferably about 1.5 to 1000 times higher, and more preferably about 2 to 100 times higher, than the outlet surface velocity during the combined reaction and sorption stages.

[0186] In one embodiment, the desorption process is enhanced by using a reduced inlet pressure at the pipe inlet to assist in the desorption of ammonia from the sorbent.

[0187] In alternative embodiments, the desorption process is further enhanced by increasing the temperature to help release absorbed ammonia from the solid absorbent. When the temperature rises above a critical value, the coordination number of absorbed moles of ammonia per mole of absorbent is reduced, thereby aiding the desorption of the resulting product. For the absorbent MnCl2, when the temperature is increased to about 300-400°C, preferably 330-390°C, and more preferably about 350-380°C, the metal halide can no longer retain the same number of moles of ammonia per mole of absorbent.

[0188] In one embodiment, a high-temperature heat transfer fluid stream, such as high-temperature oil or an equivalent, may be used from a first part of a series of processes that includes only a catalyst. In this embodiment, the first part of the process may react about 10–40% nitrogen to form ammonia at a peak temperature higher than the temperature permissible in the integrated catalyst and absorbent system. The corresponding heat transfer fluid, operated in co-current or reverse current mode in the first series process, may then flow into the desorption stage of the process. Desorption may be carried out sequentially such that the heat transfer fluid from the first reactor-only stage is supplied to a portion of the desorption tube for a time shorter than the cycle time before moving to at least a second portion of the desorption tube. All desorption tubes are sequentially regenerated to substantially cleanse the absorbed ammonia before initiating fresh reaction and sorption cycles. Substantially, purification means the removal of approximately 50–99.9999% of the absorbed ammonia before the cycle switchover, preferably approximately 50–99.999% of the ammonia, and more preferably approximately 85–99.999% of the ammonia, thereby allowing the previously cleaned desorption bed to be reused for fresh reaction and sorption.

[0189] In alternative embodiments, unreacted feedstock, such as that present from other unit operations in the system, may be used to assist desorption by increasing the sweep gas rate in whole or in part, thereby increasing the desorption rate, reducing the average residence time of desorption, thereby reducing the amount of ammonia decomposition and maximizing the rate of production of beneficial ammonia products.

[0190] In alternative embodiments, unreacted feedstock collected after a final ammonia purification step, which may include ammonia condensation or other unit operations, or unreacted feedstock collected as effluent from a combined sorbition and reaction process, or a desorption process, or a combination thereof, may be fed into a second, smaller, integrated reaction and sorbition reactor system. The second integrated reaction and sorbition system may be smaller than the first integrated reaction and sorbition system, as defined by a reduced inlet process mass flow or process volume, or a combination thereof. The inlet pressure to the second integrated reactor and sorbition system may be lower than the inlet pressure to the first integrated reactor and separation system to minimize gas compression costs while achieving a higher net overall process conversion of nitrogen and hydrogen to ammonia. The overall ammonia production rate is determined by adding the value obtained by multiplying the inlet supply rate to the first process by the conversion rate and reaction stoichiometry in the first process, the value obtained by multiplying the inlet supply rate to the second process by the conversion rate and reaction stoichiometry in the second process, and then subtracting the ammonia loss due to decomposition during desorption from that value.

[0191] This disclosure assumes that the gas collected during and after each desorption cycle contains ammonia, as well as unreacted nitrogen and unreacted hydrogen. The concentration of ammonia is expected to be substantially higher than the concentrations of unreacted nitrogen and unreacted hydrogen, which have a mole fraction of about 0.4 to 0.999999, preferably about 0.5 to 0.99999, and more preferably about 0.8 to 0.9999. The concentrated ammonia gas mixture can be desorbed at a predicted pressure lower than the inlet pressure of the first sorption and reaction cycle. The gas is expected to flow through a unit operation which may include one or more heat exchangers and / or condensation systems to capture purified liquefied ammonia for use and sale of the product.

[0192] The condensation of ammonia is expected to follow a vapor-liquid equilibrium, which is a function of both temperature and pressure. Condensation is an exothermic process, and therefore heat must be removed to allow the phase change of gaseous condensed ammonia to the liquid form. Fixed gases of nitrogen and hydrogen are non-condensable under the expected condensation process operating conditions. The plant can be thermally integrated so that the heat removed during cooling and / or ammonia condensation can be reused in the process. In one embodiment, the removed heat can be recovered to reheat the non-condensable gas, either whole or partially, before reuse in a second integrated sorption reaction system.

[0193] In one embodiment, an outlet desorption mixture gas containing ammonia, nitrogen, and hydrogen may be collected at a temperature of about 200°C to about 400°C, preferably about 300°C to about 395°C, more preferably about 330°C to about 390°C. The collected process stream may enter a heat exchanger or heat exchanger-condenser combination system so that the collected desorption stream is cooled to about -10°C to about 60°C, where condensation may occur depending on the operating pressure as phase equilibrium. For example, at about 50°C, the condensation pressure is about 20 bara; at about 20°C, the condensation pressure is around 8 bara; and at -10°C, the condensation pressure is about 3 bara.

[0194] In one embodiment, a recoverable heat exchanger transfers heat from 380°C to 25°C in a 3 bara inlet desorption stream in counterflow mode, to which a collected non-condensable gas stream discharged from a condensation unit operation flows in at approximately -10°C and is heated to approximately 300°C in an AC heat exchanger used in a second integrated reactor-sorbing system. The heat exchanger may be compact or reinforced to reduce process volume, or may include a conventional shell and tube heat exchanger.

[0195] In one embodiment, a heat exchanger and / or one or more condensers may use separate working fluids to enhance heat transfer and further cool the fluid to condensation conditions by using either a separate cooling fluid or heat transfer fluid that flows as a single phase in a closed-loop configuration or evaporates at a temperature close to the condensation temperature. This configuration allows for effective condensation by removing the heat of condensation of ammonia through closed-loop evaporation of the heat transfer fluid, so that enthalpy is transferred from the ammonia condensation stream to the evaporation working fluid stream. The energy collected in the closed-loop evaporation fluid stream or single-phase heat transfer loop can be removed to the surroundings by using a cooling tower partially or entirely. Alternatively, the energy collected in the closed-loop evaporation stream can be reused entirely or partially elsewhere in the process.

[0196] This disclosure describes how one or more reactors may be configured as shell and tube processes, and how either the shell side or the tube side of one or more reactors is filled with the catalyst and / or sorbent and / or catalyst-sorbent structure of this disclosure.

[0197] In some embodiments of the active material filling the tubes, the tubes may be operated in series or in parallel. Each individual tube within a common shell may be in either an sorption mode or a desorption mode, or each tube within the same shell may be operated within the same mode.

[0198] In some aspects of operation in the same mode, the shell may be circulated with a heat transfer fluid, and in sorption mode, the heat transfer fluid may remove heat from the tubes. This heat transfer fluid may be directed to other parts of the plant so that the heat can be used. Preferably, the heat transfer fluid may circulate to a second (or more) tubes operating in the opposite mode. For example, the heat transfer fluid may remove heat from a reactor where all tubes are in sorption mode and then circulate to a reactor where all tubes are in desorption mode. The heat transfer fluid may be in the form of a boiling liquid that can remove heat outside the reactor wall if partially evaporated.

[0199] In some embodiments, where tubes within a common shell are simultaneously in different modes, they may be configured as a bundle that provides thermal communication between them. The heat transfer fluid may allow heat to pass between individual tubes in opposite modes, or the fluid may allow heat to pass between the bundle of tubes in opposite modes. The shell may be a partitioned, shielded, or otherwise separated flow of the heat transfer fluid.

[0200] Recognizing that some catalysts require reduction at temperatures higher than those of commonly available liquid heat transfer fluids, the process of the present invention may integrate a resistor or other heating element or other method to achieve a preferred catalytic reduction temperature.

[0201] Desorption may require temperatures higher than those of the liquid heat transfer fluids typically available. Secondary heat sources can be integrated into the reactor in the form of resistance heaters, induction heaters, or similar configurations. They may also take the form of gas preheaters located upstream of the reactor unit. A secondary advantage of any such device is that it increases the rate at which the floor can heat up from a low temperature state, which is beneficial for rapidly increasing production at a constant rate from ambient or standby states.

[0202] Increasing the diameter of fixed-bed reactors and sorbents may reduce the mass of steel required for the plant, given that the thickness of the metal walls increases with increasing diameter to withstand the process temperature and pressure, but otherwise fewer pipes are needed for the same production capacity. Smaller pipe diameters may require more pipes per plant, resulting in an increase in the total metal required for fixed ammonia production capacity. Smaller pipe diameters are more responsive to thermal changes, reducing the response time of plant startup and dynamic load changes.

[0203] Heat transfer to and from larger pipe diameters for packed bed processes may respond more slowly than for smaller diameter beds with the same wall conditions or wall heat flux. The bed can be set up in stages with heat transfer interposed between the catalyst and sorbent beds. Alternatively, the bed may have multiple gas inlets along its axial length. Lower-temperature process gases can be introduced through multiple gas inlets to cool the bed and gas flow. Multiple gas inlets have the added benefit of adding molecules to the system. This is beneficial in compensating for molecules lost in the reaction and sorption to ammonia, otherwise the pressure would drop along the length of the bed. Introducing the process gas at multiple points across the bed may improve pressure stability.

[0204] While increasing the floor diameter can reduce the total plant mass of steel required for a fixed production capacity, heat transfer through a larger diameter floor is less efficient than through a smaller diameter floor with the same wall conditions. In one embodiment, the cooling medium may be contained within a packed bed comprising tubes, coils, or similar fluid chambers through which a heat transfer fluid can flow. The internal heat transfer chamber has the secondary advantage of reducing the cross-sectional area of ​​the floor, thus increasing the surface velocity of the process gas, and thus increasing the heat transfer rate between the process gas and the innovative pellets containing the catalyst and adsorbent.

[0205] In certain embodiments where a heat transfer fluid cools a reactor, this fluid may be in thermal communication with a steam generator. The steam can be combined with an electrolytic cell, particularly a solid oxide electrolytic cell. Alternatively, the fluid may be in thermal communication with any electrolytic cell for hydrogen production to provide heat, thereby improving the efficiency of the plant.

[0206] The heat transfer fluid may be implemented as a thermal energy storage mechanism, and the fluid system is adequately insulated to retain heat for hours, days, weeks, or even longer. The fluid may be discharged from the reactor shell and stored in a tank to minimize heat loss. The still-hot fluid can then be recirculated through the reactor shell to reheat the reactor bed.

[0207] This disclosure intends to use a sufficiently active ammonia synthesis catalyst such that the reaction temperature is sufficient to manipulate the active sorbent, and the active sorbent may contain a metal halide or other absorbent, adsorbent, or combination thereof, in a range in which more than 1 mole of ammonia is retained per mole of active absorbent. In the case of a MnCl2 absorbent, it is predicted that 2 moles of ammonia can be retained per mole of absorbent at temperatures of about 130°C to about 260°C.

[0208] In one embodiment, the combined reactor-absorbent system is used for a catalyst with activity equivalent to that described in Example 2 below, at a wall and inlet temperature of 320°C, with a load of 800 kg / m³. 3 Supported with an absorbent exceeding 150 kg / m³ 3 The catalyst is operated with a diameter of approximately 0.02 meters, a length of 10 meters, a pressure of 40 barg, a feed ratio of H2 to N2 of 3:1, and other catalyst pellet dimensions as listed in Table 10. With an inlet surface velocity of 1e-4 m / sec, the initial nitrogen conversion rate is approximately 95.7%, and the cycle time is approximately 50-60 hours. The number of tubes in this example is approximately 23,000 for a 4 tpd ammonia production system.

[0209] In the alternative scenario, increasing the diameter to 0.04m reduces the number of pipes to approximately 5600 for a conversion rate of approximately 96.6% under equal conditions. The temperature rise of the floor increased by approximately 1°C to 3°C within the same cycle time range. Alternatively, in this example, increasing the pipe diameter to 0.06m further reduces the number of pipes to approximately 2300 for an initial conversion rate of 96.3% and a cycle time of 50-60 hours.

[0210] (In the case of MnCl2 used as an absorbent) Operating at temperatures below approximately 330°C is generally understood to allow the absorption of more than 1 mole of ammonia per mole of active absorbent, which is desirable to enable longer cycle times. Cycle times of approximately 1 to 200 hours, preferably 4 to 100 hours, and more preferably 8 to 60 hours are desirable. A higher absorption capacity than defined by the number of moles of ammonia retained per mole of absorbent allows for longer cycle times.

[0211] At each cycle switch, unreacted nitrogen and hydrogen gases retained within the porous pellets, in the gaps between pellets, and in the headers and footers of the tube assembly are lost to the purge stream, thereby reducing the mass efficiency of the process, or they merge with the desorbed ammonia flowing into separate unit operations for purification, which may involve condensation. The unreacted feed gas after the ammonia purification step may be lost or further converted in separate reactor and / or reactor absorber systems. It is desirable to reduce the mass flow rate of this non-condensable gas stream so that further utilization or conversion can occur at smaller process volumes. Achieving longer cycle times may be advantageous for the combined reaction-sorbation system.

[0212] This disclosure describes how the reactor-absorber system of the present invention may include two or more sections connected in series, where the flow is intended to move sequentially from a first section to a subsequent section during a process cycle before the feed is switched to a parallel reaction and absorption system. In one embodiment, the first reactor section includes a catalyst, and the second reactor section includes a catalyst and an absorbent in close contact within a catalyst-sorbent structural component such as a pellet.

[0213] In an alternative embodiment, two reactor sections operated in series each contain a catalyst and an adsorbent in close contact within a catalyst-adsorbent structural component such as a pellet. The first and second sections may have different tube diameters and lengths, as well as different catalyst and / or absorbent loading densities within the catalyst-adsorbent structural component.

[0214] In one embodiment, the reactor-absorber system of the present invention may have three series-connected sections, wherein the first section contains a catalyst, and the second and third sections contain a catalyst and an absorbent in close contact within a catalyst-sorbent structural component such as a pellet. The number of tubes in each section may be the same or different between sections. In one embodiment, the first section may have at least 10% fewer tubes than the second section, preferably about 10-95% fewer tubes, and more preferably about 20-90% fewer tubes.

[0215] The number of tubes in two series-connected sections containing a catalyst and an adsorbent in close contact may be approximately 1 to 100, or approximately 100 to 1, from the first series section to the second series section.

[0216] In one embodiment, the hybrid two-compartment series integrated reactor and absorber may be operated in a first section of 8 m in length and a second section of 2 m in length, where the catalyst loading density for a catalyst having activity equivalent to that described in Example 3 (hereinafter) is 125 kg / m³ in the first section. 3 , in the second category 150 kg / m 3Both sections were operated with 40 barg, and the absorbent density was 800 kg / m³. 3 The pipe diameter of the first section is 0.04 m, the pipe diameter of the second section is 0.03 m, the pipe wall temperature of the first section is 305°C, and the pipe wall temperature of the second section is 310°C. The net reactor volume of the two-section series integrated reactor and absorber can maintain a peak internal floor temperature below 330°C so that the absorbent can capture or absorb approximately 1 mole of ammonia per mole of absorbent. The net volume of the two-section series reactor is less than 50% of the volume of a pipe of equivalent length with a diameter of 0.03 m. The outlet of the first section flows into the second section, which has the same mass flow and composition. In this case, a higher feed rate can be used in the two-section reactor by adjusting the catalyst density, pipe diameter and length, and temperature so that the rise in internal floor temperature remains below the temperature transition point where the absorbent can no longer hold more than 1 mole of ammonia per mole of absorbent.

[0217] It is generally understood that specific hybrid designs can be optimized for each of two or more segments, depending on catalytic activity and temperature switching point.

[0218] In alternative embodiments, a reactor system having two or more series sections may be operated such that the desorption process occurs in a manner in which a first section or more sections are first desorbed with respect to a plurality of tubes operated in parallel, and then a second integrated reactor and absorber section or more sections are desorbed in parallel. In one embodiment, the desorption of the second section may be initiated before, during, or after the feed is switched to the first section. In the process of the present invention, in which the desorption of the second section is initiated after the feed has been returned to the first section in series, unreacted nitrogen and hydrogen leaving the first section may function as desorption aids or scavenging gases to the second section. The temperature of the second section may be increased to aid desorption by using a heat transfer chamber surrounding one or more tubes operated in parallel, thereby the temperature of the heat transfer fluid is higher than the temperature of the heat transfer fluid in the first section due to the reuse of heat generated from the combined catalyst and absorber tubes or tube bundles elsewhere in the catalyst-only reactor and / or system.

[0219] In alternative embodiments, the tube may include the use of electrical resistance or induction heating to increase the wall temperature to assist in the desorption of ammonia. It is desirable to enable digital control of temperature so that there may be two or more thermal zones achieved by electrical means on any section of the process of the present invention.

[0220] This disclosure envisions an integrated reactor and absorber system operated in series sections such that the conversion rates on the first section and subsequent sections are maintained at approximately 30–90%, and as a result, a relatively high gas surface velocity can be used in the first section to improve heat transfer while achieving a lower conversion rate. The convective component of pellet heat transfer and associated gas heat transfer to the walls increases at higher surface velocities. Improved floor heat transfer allows for the use of larger pipe diameters, reducing the total number of pipes in the plant with equal production capacity, along with the associated costs of metal, flanges, piping, and assembly. The unreacted feed from the first section, with a conversion rate within a defined range, can then be sent to separate separation steps, including absorption, adsorption, or a combination thereof, to separate the generated ammonia. The unreacted feed can then flow into a second reactor and absorber system, where the reaction can continue for a higher net or overall feed utilization rate.

[0221] The pressure in the second integrated reactor and absorbent system may be the same as, lower than, or higher than that in the first integrated system. If the pressure in the second system is lower than that in the first system, little to no gas compression is required. It is desirable to operate the systems in series, with each of the two or more systems operating at a lower pressure, to avoid or minimize gas recompression and associated plant complexity, unit operation, and costs.

[0222] In one embodiment, the three reactor-absorbent sections are operated in series with a catalyst of equivalent activity to that described in Example 3 (and below), as shown in Table 1, and the integrated reactor and absorber of the three sections can minimize volume and metal weight by operating each stage at a higher rate, lower conversion rate, and larger tube diameter.

[0223] [Table 1]

[0224] Reductions in the number of tubes, reactor-absorber volume, and / or metal weight can be achieved by operating integrated reactor-absorbers in a section, thereby increasing the diameter and rate and potentially lowering the section conversion rate; however, when multiple sections are operated in series, the overall net conversion rate of nitrogen may remain high.

[0225] In one alternative embodiment, a continuous, segmented reactor may be regenerated to desorb the absorbed ammonia in a continuous manner, so that at least one segment desorbs during sorption in the first segment.

[0226] The following embodiments provide further disclosure of the present invention. However, it should be understood that the intention is not to limit the claimed invention to the specific embodiments described in the following embodiments, but rather to provide them merely as illustrative examples. [Examples]

[0227] [Example 1] The production of ammonia using highly active cobalt-based catalysts has been described by Gao and colleagues. This 2017 publication provides data on the low-temperature production of ammonia from nitrogen and hydrogen for cobalt, supported by carbon nanotubes and tested at 10 bara. Gao reported the reaction orders of hydrogen and ammonia as 0.58 and -1.27, respectively. A series of power-law reaction rates were fitted to the available data, as published by Gao and colleagues. The form of the reaction rate equations included the standard equilibrium form for thermodynamic consistency. Values ​​of the equilibrium constants were provided by Sehested and coworkers, 1999.

[0228] The rate of ammonia production is described below, and the nitrogen consumption rate is half of this rate.

number

[0229] Figure 6 shows the ammonia production rate per hour per gram of catalyst reported by Gao in 2017. The solid black circles are the reported values, and the open squares are the kinetic fitting values of this reaction. The reaction rate parameters were fitted using an automatic genetic algorithm that links DAKOTA with a commercially available DETCHEM PBR reactor code to optimize the parameters for minimizing the residuals. The fitted values of the reaction rate of the Co-based catalyst shown in Gao's literature were evaluated in a recycle reactor to serve as a benchmark for the predicted performance of the comparison system with the present invention. Gao's data and the modeled reaction rate fitting data are shown in Figure 6. The ammonia production rate per hour per gram of catalyst reported by Gao in 2017 is shown as solid black circles, and the kinetic fitting modeled data of this reaction are shown as open squares.

[0230]

Table 2

[0231] Regarding the table summarized in Table 2, those results are based on the input conditions summarized in Table 3. The temperature refers to a constant wall temperature and inlet gas temperature. The velocity is superficial as calculated by the inlet temperature and pressure.

[0232]

Table 3

[0233] Note that the inlet molar fraction of ammonia is set to a low but non - zero value to avoid mathematical singularities in the form of a rate equation that includes a negative order with respect to the ammonia partial pressure. The sensitivity of this value is such that a sufficiently low value that does not affect the results is chosen, and the range of the inlet molar fraction is from 1e - 5 to 1e - 7.

[0234] The recycle ratio (RR) is calculated and performance - evaluated for a recycle system in which all ammonia is removed in a series of downstream ammonia separation steps. Ammonia can be separated by absorption, adsorption, condensation, or other methods.

[0235] The feed rate to the reactor inlet is calculated as the fresh feed rate based on 100% mass conversion to achieve the capacity indicated on the plant and is multiplied by the defined recycle ratio. The mass balance is checked to ensure that the ammonia produced from the reactor for each test condition is equal to the production capacity of the target plant. In an industrial system, an over - design factor in the range of 3 - 15% is predicted, at least 1%, or preferably 3 - 10%, to account for system losses in downstream unit operations.

[0236] The design calculations are based on an ammonia production plant capacity of 4 tpd. A similar series of analyses were performed using reaction - rate theory fitting for Co - based catalysts from Gao 2017 for a non - isothermal reactor that includes pellet - internal pore diffusion of reactants and products as presented in Table 4. The same reactor assumptions as described in Table 3 were used.

[0237]

Table 4

[0238] Using reaction rates as fittings from Gao 2017, the expected performance improvements when used in the reactor of the present invention, in which the catalyst and absorbent are closely integrated, were evaluated. Typical catalyst loading densities in commercially available pelletizing reactors are approximately 800–2000 kg / m³. 3 This range is possible. Some catalysts may be outside this range, as in the case of high iron-based catalysts. Lower catalyst densities can be achieved by the system of the present invention through the structures of the intermixed, single-coated, or double-coated embodiments.

[0239] For highly active catalysts like those described in Gao 2017, it is difficult to operate them with high volumetric catalyst loading and a high conversion rate per pass while maintaining the peak temperature below approximately 330°C, so that ammonia continues to be absorbed in a 1:1 molar ratio with the metal halide. The results are shown in Table 5, and the temperature was at least 300 kg / m³. 3 Before the absorbent is filled into the reactor along with ammonia, accompanied by the catalyst supporting the above-mentioned amounts, the ammonia absorption limit defined by the metal halide may be exceeded in the initial part of the cycle under various reactor configurations and conditions. For MnCl2, a peak temperature of approximately 330°C or approximately 320°C to 350°C is predicted.

[0240] [Table 5]

[0241] Approximately 100kg / m 3 Table 1.5 shows low levels of catalyst loading below a certain threshold. For highly active catalysts, the acceptable operating window can be defined by reducing the amount of catalyst loaded in the system of the present invention.

[0242] [Table 6]

[0243] The loading density of the absorbent is 800 kg / m³, which has the properties of MnCl2. 3 It is evaluated in this way. At room temperature, the fully densified metal halide has a density of approximately 2900 kg / m³. 3 It has a density exceeding [a certain value]. The halogenated absorbent is understood to be contained in or supported within a porous support and pelletized. The absorbent is arranged continuously, discontinuously, or continuously-discontinuously within the support. In this embodiment, the porous support is included up to about 25% to 30% of the reactor volume. In some embodiments, the absorbent occupies a portion of the reactor, and the absorber volume is about 10 to 80%, preferably in the range of about 20 to 40%. Within this range, the absorbent can change from a more discontinuous structure to a continuous-discontinuous structure. In this structural transition to a continuous region containing the bound absorbent, the effect of osmosis theory can enhance mass transfer that moves ammonia within the passage of the bound absorbent. In this embodiment, the sorption capacity is enhanced between the reaction and sorption portions of the cycle described above, and desorption is enhanced when no feed is flowing into the specific reactor vessel or tube containing the sorbent and catalyst, as described by the present invention.

[0244] In one embodiment, the volume increase during ammonia sorption is limited to a localized area, and the overall macroscopic pellet dimensions remain substantially unchanged or do not swell. In an alternative embodiment, the volume increase in the localized area acts to bind other fragmented areas of the absorbent, thereby increasing the mass transfer rate and, by extension, increasing the effective volume or utilization rate of the absorbent by allowing ammonia to move more easily to areas with lower packing capacity.

[0245] For test conditions that exceeded a conversion rate of 95% per pass, maintained a peak temperature below approximately 330°C, and reduced absorption capacity, the results from Table 1.5, sorted from high to low by conversion rate and highlighted, are shown in Table 1.6.

[0246] [Table 7]

[0247] By reducing the amount of catalyst supported in the reactor for a highly active catalyst, the reactor volume increases so that ammonia can continue to be absorbed, enabling a high conversion rate per pass while maintaining a peak temperature below about 330°C. Note that in generally reported simulations, ammonia can continue to be absorbed at temperatures above 330°C. For an exemplary MnCl2 metal halide, the absorption capacity of NH3 decreases to about 0.5 mol / mol at about 330°C to about 370°C and then substantially decreases to zero at higher temperatures.

[0248] For the MnCl2 absorbent, it is desirable to maintain the temperature of the integrated bed below about 330°C, or about 320°C to about 370°C, or more preferably below about 300°C to about 320°C. It is desirable to identify absorbents with higher ammonia capacity and absorbents that retain absorption capacity at higher temperatures and / or lower pressures.

[0249] The cycle time required to switch the feed from the first integrated reactor to the second parallel integrated reactor depends on the holding capacity of the selected metal halide absorbent and its amount in the integrated reactor and absorbent bed.

[0250] The absorption rate on MnCl2 metal halide is estimated from Smith and Torrente-Murciano 2021 as a function of the absorbent loading density. The absorption rate is substantially faster than the initial reaction rate (peak value) across the range of expected operating conditions. In an exemplary case where the amount of catalyst is minimized and the amount of absorbent is maximized to enable a longer cycle time, the absorption rate is higher than the reaction rate. Due to mass conservation, the local ammonia absorption rate when the catalyst and absorbent are in close contact cannot exceed the ammonia formation rate and thus matches.

[0251] For a design space in which a MnCl2 metal halide absorbs 1 mole of NH3 per mole of absorbent, the theoretical absorption capacity of the material is approximately 0.1353 g of NH3 / g of active absorbent. Furthermore, additional ammonia may be adsorbed onto the surface of the MnCl2 or the support or catalyst material. The metal halide salt may be supported in a porous structure and pelletized, or pelletized in such a way that it maintains sufficient internal porosity for gas-phase diffusion of ammonia, thereby allowing access to the metal halide material while avoiding solid-state diffusion. The diffusion process within an exemplary pressed pellet may be Knudsen diffusion of molecules, as determined by the size of the internal pores and the mean free path of ammonia.

[0252] The expected loading density for metal halide absorbents or other sorbent materials is approximately 100 kg / m³. 3 ~2000kg / m 3 The range is preferably about 300 to 1500 kg / m 3 More preferably, within the range of approximately 500-1200 kg / m 3 This range is possible. The cycle time and performance of the absorption and reaction system of the present invention are shown in Table 1.7.

[0253] [Table 8]

[0254] The extremely high loading capacity of highly active catalysts (300 kg / m³) as described by Gao et al. in 2017. 3 For cases with a value exceeding 500 kg / m³, the range is approximately 1500 kg / m³. 3The cycle time for the absorbent-loaded material is less than approximately one hour. One trade-off analysis suggests that while very fast cycle times (less than one hour) minimize the overall reactor volume, the trade-off is a less robust system that requires more frequent cycles. In each cycle, unreacted feedstock is lost within the pores of the pellets, within the gaps between pellets in the packed bed, and within the void volumes in the headers and footers connecting the tubes, which would reduce the overall mass efficiency of the process of the present invention.

[0255] The expected design is generally understood to maximize the absorbent load to increase cycle time, thereby reducing the frequency of floor cycles and regeneration. The feed to the reactor is circulated from the absorber of the parallel, integrated reactor before the floor capacity reaches the theoretical load. The process feed is recycled when the floor reaches a rate in the range of about 0.1 to about 0.9 of the theoretical floor load, preferably in the range of about 0.2 to about 0.85, more preferably in the range of about 0.3 to about 0.8.

[0256] Furthermore, it is anticipated that increasing the pipe diameter will be explored as a more optimal configuration, thereby reducing the number of pipes required, such as flanges, manifolds, and connecting pipes, while simultaneously reducing the weight of the associated metals to be procured. The assumed wall thickness of 3 mm will be optimized for the duty cycle, as well as the associated pressure, chemical, and thermal stresses.

[0257] [Example 2] The multistage system may be configured such that the first stage involves only the reaction without separation. The product from the first stage flows to one or more integrated reaction-sorbing reactor stages, as described in the present invention.

[0258] In an alternative embodiment, the process is configured to include a first catalyst-only stage to produce a portion of the target plant capacity of ammonia. This first stage has a transformation limited by equilibrium according to thermodynamics. In a preferred embodiment, there is no recycling of unreacted feed to this first stage to minimize the need for process compression and heating. This process continuously produces about 10–40% of the plant capacity in the first stage containing only the catalyst.

[0259] As shown in Table 8, the reaction-only process has lower heat release than the combined sorption and reaction system, and allows for larger predicted tube diameters and higher allowable peak temperatures to minimize the overall reactor volume, number of tubes, and metal weight in the first stage. The analysis is completed at reaction rates that fit the data provided in Gao 2017, as described in Example 1.

[0260] [Table 9]

[0261] In the first stage, in an example case with a conversion rate of 25% per pass, approximately 5-7m 3 With a reactor volume and an estimated metal weight of approximately 0.5–2 tons, it is possible to produce approximately 1 ton of ammonia per day. These metrics represent a reduction of approximately 10–50% in reactor volume and approximately 10–98% in metal weight in the first part of the process for a process containing only absorbent and catalyst in all reaction vessels. Due to remaining equilibrium limitations, this hybrid process can only produce approximately 10–40% of the total plant capacity in the first stage, and the unconverted feedstock is not wasted as it is supplied to the second stage, which contains pellets of the catalyst and absorbent of the present invention in close contact with the materials.

[0262] For the comparison of fully integrated reactor and absorber systems shown in Table 1.6, the reactor-absorber tube volume for a conversion rate of over approximately 95% is approximately 55-80 m³. 3 This requires a hybrid embodiment, which is a means of reducing the overall reactor volume and the required metal weight. For a fully integrated reactor and absorber system, the first tonnage of ammonia per day is approximately 13-20 m³, with a wall thickness of 3 mm and a metal weight of 130-350 tons, representing about 25% of the total volume. 3 This will be necessary. Generally, it is understood that the final design may require an optimal wall thickness, which can range from about 1 to 5 mm, for the required duty cycle and mechanical robustness. Accordingly, the weight and weight reduction benefits for the hybrid process of the present invention will vary.

[0263] In the second stage of this hybrid process, the generated ammonia can be supplied to an integrated reactor and absorber. The ammonia generated in the first stage is captured on the absorbent, and the reaction continues to proceed at a higher conversion rate because the thermodynamic limit of the reverse ammonia decomposition reaction is restricted by localized and close product removal.

[0264] The hybrid process of the present invention can be further advantageous by using three or more steps in series. In the first step, only the reaction proceeds, allowing the production of approximately 10-40% of the plant capacity. In the second step, the ratio of absorbent to catalyst density is higher than that of the third step in series. The ammonia produced in the first step is preferentially absorbed in the second step. The third step continues to react most of the nitrogen with hydrogen, achieving a net process conversion rate of approximately 80-99.9%, preferably 90-99%, and more preferably about 95-98%.

[0265] In an alternative embodiment to this hybrid process, a first highly active catalyst formulation, such as the catalyst described by Gao in 2017, may be used in the first step, and a second catalyst formulation, which may be less active, may be used in at least the second step of the hybrid process. The active catalyst material may be the same, different, or a combination of the two in at least each of the first and second steps.

[0266] [Example 3] Alternative catalysts based on formulations containing iron and cobalt are described in Smith and Torrente-Murciano (2022) for a commercially available composite Fe-Co catalyst available from Johnson Matthey under the trade name Katalco 74-1.

[0267] The data is approximately 1e-6m 3 The Smith paper presents the low catalyst load (0.2g) and high catalyst load (1.5g) contained within the reactor volume. In both cases, the inert material is mixed in at the low load to maintain the same packed-bed reactor volume. The higher bed load (1.5g) is approximately 1492 kg / m³. 3 Catalysts with a loading density of (approximately 190 kg / m³) yielded lower loading volumes. 3 This produced a conversion rate approximately twice that of 0.2g for a loading density of 0.2g. The conventional Temkin-Pyzhev power law form of the reaction rate equation did not adequately capture the data. We evaluated an alternative form of the reaction rate equation based on competitive Langmuir adsorption at surface sites, which takes the form of LHHW (Langmuir-Hinshelwood-Hougen-Watson) reaction kinetics, and fitted optimized reaction rate parameters by regression on the dataset described in the publication.

[0268] The reaction rate for ammonia formation is shown below, in which the nitrogen consumption rate is half the ammonia formation rate and follows the stoichiometry of the reaction.

number

Number

[0269] As presented in the paper by Smith 2022, an equilibrium comparison between the reaction rate prediction and the data is shown in Fig. 7. The complexity of the catalytic reaction rate can be seen in the data, and an increase in the catalyst by 7.5 times per reactor volume doubles the nitrogen conversion rate approximately. The influence of the reaction equilibrium with the reverse reaction (ammonia decomposition), combined with the influence of ammonia adsorption on the catalyst sites, strongly inhibits this Fe-Co catalyst. Although the reaction rate equation of the present invention does not fully conform to complex chemical reactions, it is generally understood that it reflects a reasonable approximation of the current and predicted performance when scaling the catalyst with an absorbent in the present invention.

[0270] The reaction activation energy is 140.06 kJ / mol. The heats of adsorption of NH3, H2, and N2 are -32.73, 23.85, and -24.41 in kJ / mol units, respectively.

[0271] The pre-exponential term of the reaction is 8.35e29 in mol / gram / min / bar units, and the pre-exponential terms of the adsorption of NH3, H2, and N2 are 3.17e 16 , 1.994e 3 , and 7.25e 11Catalysts tested in small granular form have, in most cases, shown neither internal pore diffusion limitation nor substantial temperature rise, as defined by a temperature of less than approximately 1°C.

[0272] The adsorption constants for each species are shown in Figure 8 as a function of reaction temperature over the range of test data. The adsorption constant for hydrogen decreases with temperature, while the adsorption constants for ammonia and nitrogen increase with temperature.

[0273] Since the denominator term is squared in the rate equation, the effect of inhibiting ammonia from surface adsorption on the catalyst site significantly limits the overall rate of ammonia synthesis. Coupled with the thermodynamic limit on the overall conversion rate, this limitation of surface adsorption limits the overall rate of ammonia formation. A refined rate equation is expected to achieve a close balance with experimental reactor data. Further evaluation of the effects of scaling using cobalt and iron-containing catalysts is needed to understand the effect of these catalysts in recycle reactor configurations, which are operated with integrated catalysts and absorbents maintained in close contact, as described in this invention.

[0274] Table 9 shows the effects of manipulating the catalyst in the recycling reactor system, as described by Smith in 2022. The reactor operates upstream of a separate unit operation including a separator, and the generated ammonia is removed by absorption, adsorption, condensation, or other means. Unreacted feedstock is recompressed and recycled to the first reactor. Input conditions are listed in Table 10.

[0275] [Table 10]

[0276] [Table 11]

[0277] [Table 12]

[0278] The overall catalytic activity of Katalco 74-1 per unit weight is lower than that of the catalyst with only cobalt dispersed on carbon nanotubes, as explained by Gao 2017 and as described in Example 1. 2200 kg / m in Example 1 3 versus 1200 kg / m 3 At the higher catalyst loading density in this case of, for this combined iron-cobalt catalyst, the recycle reactor volume becomes larger.

[0279] [Example 4] Integrated pellets containing MnCl2 metal halide for ammonia absorption and an ammonia synthesis catalyst containing iron and cobalt based on the reaction kinetics that match the data of Katalco 74-1 described in Example 3 were investigated for their performance under integrated operation, various conditions and loadings.

[0280]

Table 13

[0281] Using the reaction rate for Katalco 74-1, the internal bed temperature as a function of length was evaluated for a plurality of catalyst loading densities and shown in Table 12. For cases other than 100 kg / m 3 the conversion per pass exceeds 95% and the specific results are shown in Table 12. The non-listed packed bed parameters in Table 12 are as described in Table 10.

[0282] For catalyst loadings of about 100 - 300 kg / m 3 as shown in Figure 9 for the selected reactor dimensions and conditions, the bed temperature is less than about 330 °C.

[0283] It is understood that when the internal temperature exceeds approximately 370°C, the absorbent bed will substantially cease to retain or absorb ammonia, and when the bed temperature exceeds approximately 800°C, the catalyst may begin to sinter due to permanent inactivation. The results in Figure 10 show that higher catalytic activity or loading (approximately 400-1200 kg / m³) causes the hot spot to move towards the front of the bed. 3 This is an example to illustrate the effect of ). The results shown in Figures 9 and 10 are early snapshots of the thermal profile of the floor over cycle time before the absorbent begins to be substantially filled, so that the supply is circulated to fresh floor, when the absorbent is filled to 10-90% of its theoretical capacity (preferably about 30-85% of the theoretical absorbent capacity, and more preferably about 40-70%).

[0284] The thermal profile of the combined absorbent and catalyst of this invention, maintained in close contact, depends on the specific catalyst formulation as well as the weight loads of the catalyst and absorbent. While more active catalyst yields higher exothermic effects, it is understood that this exothermic effect can be controlled by reducing the amount of catalyst loaded in the reactor volume, adjusting the tube diameter, increasing the surface velocity, and optimizing the reactor temperature at one or more locations along the reactor length.

[0285] [Example 5] Ammonia generated and absorbed during the first cycle is removed or desorbed during the second cycle, while minimizing back-reactions that decompose or reconvert ammonia into nitrogen and hydrogen feed gases. The effects of ammonia decomposition on the catalysts described in Example 3 were evaluated and summarized in Table 13. Physical properties of catalysts and absorbents, where not currently defined, are shown in Table 14.

[0286] Increasing the surface velocity during desorption has been shown to reduce the mean residence time, thereby reducing the loss of ammonia to catalytic cracking. The cracking reaction is limited by the reduced residence time, even at high temperatures that may be required to affect the coordination number, so that the sorbed ammonia is released into the gas stream and removed from the system.

[0287] [Table 14]

[0288] [Table 15]

[0289] In integrated catalyst and absorbent systems, minimizing the catalyst weight load and / or exposed surface area within the reactor helps minimize the reverse ammonia decomposition reaction during desorption. (Table 5.1, 800 kg / m²) 3 As shown, with high catalyst loads, lower gas rates result in longer average residence times in the reactor and higher desorption temperatures, which decompose more than 20% of the absorbed ammonia. The average residence time is defined by dividing the reactor length by the surface rate and multiplying by the porosity.

[0290] High catalyst load capacity (800 kg / m³) 3 Under the same conditions, with a high desorption temperature (380°C) and a higher surface velocity of 1 m / s, ammonia decomposition is reduced to less than 0.5%.

[0291] Low catalyst load (for example, 150 kg / m³) 3 In these cases, the ammonia decomposition conversion rate decreases to less than 5% at 0.01 m / s and less than 0.05% at 1 m / s. In these cases, the average residence time during desorption is approximately 400 seconds and 4 seconds, respectively.

[0292] With respect to the combined catalyst and absorbent of the present invention, the ammonia decomposition conversion rate is less than about 5%, about 0.001% to 5%, preferably about 0.002% to 4%, and more preferably about 0.002% to 0.5%.

[0293] [Example 6] Catalyst-adsorbent structures in the form of 6 mm pellets were fabricated. Fe / Co-based catalyst powder was synthesized according to the formulation of Kai et al., 2022. Separately, MnCl2 was impregnated in SiO2Aerosil by single-step aqueous impregnation in 25 wt% SiO2. The resulting powder was mixed in an adsorbent:catalyst mass ratio of 10:1, compressed (pressed) with a binder to form pellets, and calcined in air at 250°C. The resulting pellets had an average diameter of 6.1 mm and a crushing strength of 3.7 lbf. The pellets were epoxy-mounted and polished down to the core for SEM / EDS imaging. SEM images of the resulting pellets are shown in Figure 12A. EDS images of the resulting pellets are shown in Figure 12B. SEM / ED cross-sectional images show close contact between the catalyst and adsorbent, and homogeneous dispersion of the catalyst and adsorbent throughout the material, according to a certain preferred embodiment of this disclosure.

[0294] [Example 7] Zeolite 13X powder was modified to incorporate active catalysts and sorbent metals in their Na forms. 0.025 molars of metal salts (Mg(NO3)2.xH2O, Mn(NO3)2.xH2O, Cu(NO3)2.xH2O, Co(NO3)2.xH2O and Zn(NO3)2.xH2O) were dissolved in 50 ml of deionized water and added to 2 g of Sigma Aldrich Na13X, H-ZSM5, or Na-4A. The solution was stirred on a stirring plate at room temperature for 24 hours. After completion, the suspension was treated by centrifugation at 2000 rpm for 10 minutes, the supernatant was decanted, replaced with fresh DI (deionized) water, and centrifuged again. This washing process was repeated a total of four times, and the resulting solid was dried overnight by heating under vacuum at 200°C. The sample was dissolved by microwave-assisted acid digestion and characterized by ICP-OES. Subsequently, the average concentration of the relevant metal was determined, as reported in Table 15, demonstrating successful exchange to the desired metal ions.

[0295] [Table 16A]

[0296] [Table 16B]

[0297] [Example 8] 20–50 mg of the selected ion-exchange zeolite described in Example 7 was placed in TA Instruments Q50 GA. Adsorbed water was removed by heating to 500°C for 10 minutes under 100 sccm of N2. NH3 gas was introduced at 40 sccm, and the N2 was reduced to 60 sccm. This corresponds to t=0 min on the graph in Figure 13. The temperature was then gradually lowered to 300°C while maintaining a constant gas flow. The weight increase shown in Figure 13 is due to ammonia adsorption, and the relative weight increase between zeolites with varying metal loads changes with temperature, indicating that ion exchange alters the ammonia adsorption properties.

[0298] [Example 9] This embodiment demonstrates a catalyst-adsorbent structure in a system consisting of a bed of Ru-based catalyst and metal-exchange zeolite adsorbent in simulated close contact, thereby mixing small powders in approximately close contact within the packing bed by providing a large surface area of ​​catalyst and adsorbent in close contact, rather than forming a co-pressed pellet of catalyst and adsorbent. Ru-based catalysts are generally accepted to perform well at low temperatures compared to standard Fe-based catalysts. Low-temperature operation is preferable for the catalyst-adsorbent system of the present invention to achieve sufficient adsorption during reaction conditions. Zeolite adsorbents also offer advantages such as thermal stability and prevention of catalyst toxicity events.

[0299] The catalyst and adsorbent for this experiment were prepared as follows: 5% by weight Ru and 10% by weight Cs supported on CeO2 were prepared by conventional initial wetting impregnation. CeO2 support powder was prepared by calcining cerium(III) nitrate hexahydrate (Sigma, 99%) in a muffle furnace at 350°C for 2 hours. The obtained CeO2 powder was then treated with ruthenium nitrate nitrosyl (0.3 cc / gCeO2 The samples were impregnated with ) and dried in a convection oven at 80°C for 12 hours, then impregnated with a cesium carbonate solution. The resulting samples were dried in a convection oven at 80°C and sieved through a 40-100 mesh sieve before use. Magnesium-exchanged ZSM-5 (Si / Al2=23, Zeolyst) was prepared at room temperature by ion exchange by suspending 2 grams of ZSM-5 powder in an aqueous solution and adding 6.41 grams of magnesium nitrate. The solution was stirred overnight (for about 12 hours), then separated by centrifugation, washed three times with deionized water, and dried in a muffle oven at 100°C for 12 hours.

[0300] Approximately 1 gram and 0.4 grams of catalyst and adsorbent powders, respectively, were placed into the reaction chamber. The catalyst reaction chamber measured 2.75 inches (7 cm) in length, 0.205 inches (0.521 cm) in inner diameter, 0.375 inches (0.953 cm) in outer diameter, and 1.46 cm in diameter. 3 The floor volume was that of a SiC preheating zone floor exceeding 1 / 4 inch (0.635 cm). The floor was pre-reduced in 50 sccm of high-purity H2 at 400°C for 12 hours, and then cooled to the reaction temperature of 340°C at 10°C / min. The gas line was then switched to bypass the reactor, and a 3:1 H2:N2 (moles or volume) feed gas was introduced at a total flow rate of 70 or 20 sccm (measured under reference conditions of 273 K and 1 atm). The pressure was configured by a downstream dome-supported BPR (Equilibar) until a setpoint pressure in the range of 22–35 barg was reached. A downstream online mass spectrometer monitored the composition of the outlet gas flow, and a mass flowmeter monitored the total gas flow rate at the outlet. The reaction commenced when the gas line was switched from the bypass to flow into the reaction chamber. Adsorbent regeneration was performed by pressure swinging from the reaction pressure to 1 bara at 300°C, with a 60-80 sccm N2 flow acting as the scavenging gas.

[0301] Figure 14 shows the results of these experiments, which result in increased rates of NH3 adsorption and NH3 production in the integrated catalyst-sorbent structure of this disclosure. The approximately 0.3-hour delay in NH3 elution represented a breakthrough period caused by NH3 adsorption on ZSM-5 during the initial synthesis-adsorption period. After the reaction was carried out for 4 hours, it was terminated by shutting off the H2 gas flow. After the outlet NH3 and H2 signals stabilized to negligible levels on the downstream mass spectrometer, the absolute pressure to the downstream dome-supported BPR was gradually reduced by releasing the applied pressure. The mass flow rate was continuously monitored with a downstream mass flowmeter to measure the outlet molar flow rate during desorption. Approximately 0.0053 moles of intact NH3 eluted from the system over a reduced pressure range of 36–1 bara. From the amount of desorbed NH3, the N2 conversion rate during the adsorption period was estimated to be approximately 33%. This is the effect of the catalyst alone, compared to 12% during steady-state operation after the adsorbent was loaded to capacity.

[0302] As shown in Figure 14, all ammonia produced is initially adsorbed by the sorbent (zeolite). The jump in NH3 pressure occurs when the sorbent is nearly saturated. This indicates quantified performance, resulting in a better production rate when the active catalyst is in close, simulated contact with the sorbent. The rate of NH3 formation on the catalyst is slower than the adsorption rate.

[0303] [Example 12] This embodiment relates to the pre-reduction of a catalyst. The pre-reduction was performed using a material with a length of 2 inches (5.08 cm), an inner diameter of 0.205 inches (0.521 cm), an outer diameter of 0.375 inches (0.953 cm), and 1.06 cm. 3The reduction was carried out in situ in a conventional packed-bed flow reactor with a bed volume of over 1 inch (2.54-cm) of SiC preheating zone. Reduction was performed at constant temperatures of 350°C, 400°C, and 500°C maintained by an isothermal furnace, with a high-purity H2 stream of 50 sccm (measured under reference conditions of 273 K and 1 atm) and 1 bara. Pre-reduction was performed at 350°C and 400°C for 48 hours, and pre-reduction at 500°C for 12 hours. In separate cases, the catalyst was exsituated in a tubular furnace at 500°C using a high-purity H2 stream of 50 sccm, and then transferred to the reaction chamber after being exposed to ambient air during the transition. Unreduced Katalco 74-1R was placed directly into the reaction chamber and heated to the reaction temperature in 50 sccm of N2.

[0304] In-situ pre-reduction was immediately followed by an NH3 synthesis test in the same reactor. Approximately 2 g of catalyst was crushed, sieved through a 40-100 mesh sieve, and placed in the reaction chamber. During in-situ pre-reduction, the reaction chamber temperature was gradually increased at 5°C / min to less than 50 sccmH2. The catalyst was then cooled to the reaction temperature of 300°C at 10°C / min under at least 50 sccmH2. The gas flow line was then switched to bypass the reactor, and a flow ratio of 3:1 H2:N2 was introduced to match the stoichiometry of the NH3 synthesis reaction. The gas pressure was slowly built up using a dome-supported BPR (Equilibar). Once the system was equilibrated to the setpoint pressure and flow rate, the reaction was started by switching the gas line configuration back to a packed-bed flow reactor. The effluent gas stream composition was monitored by online mass spectrometry.

[0305] Table 16 summarizes the effect of pre-reduction on catalytic activity. The data in Table 16 are also shown graphically in Figure 15. Generally, in-situ pre-reduction at 500°C for 12 hours did not result in any loss of NH3 synthesis activity compared to reduction at 400°C for 48 hours. Pre-reduction at 350°C for 48 hours resulted in lower catalytic activity, with the NH3 synthesis rate decreasing by approximately 1.4 times compared to higher reduction temperatures. Katalco 74-1R reduced in excitu showed a significant activity loss of approximately 7.5 times compared to reduction at 400°C. Unreduced Katalco 74-1R was present in a volume of 1.2 μmol g. 触媒 -1 minutes -1 It reached 0.2 μmol g 触媒 -1 minutes -1 A small peak of activity was observed during ToS less than 5 hours, decaying to a low steady-state rate. Therefore, pre-reduction strongly affects the NH3 activity of the Fe-Co composite catalyst.

[0306] [Table 17]

[0307] [Example 13] Pre-ammonium treatment of metal halide sorbent The effect of pre-ammonium treatment of a halide-based absorbent (MnCl2) was investigated in a mixed catalyst-absorbent reaction bed system. The apparatus was the same as in Example 12. According to the literature, the heat treatment of metal halides can result in hydrolysis according to the reaction scheme shown in the following equation (3). MX2·H2O(s)=M(OH)X(s)+HX(g) (3) Pre-ammonium treatment can mitigate decomposition at high temperatures by exchanging absorbed H2O for NH3. Such treatment prevents the loss of gaseous halides during temperature increases in closely related catalyst-adsorbent structures, as provided in this disclosure, reducing sorbent volume and potential catalyst toxicity from Cl or X halides.

[0308] The effect of pre-ammonium treatment is demonstrated by 0.4 grams of MnCl2·(H2O) x The reaction was investigated using a mixed bed of 1 gram of Katalco 74-1R catalyst containing an absorbent. Prior to ammoniumization, the absorbent was pretreated in a tubular furnace at 350°C for 12 hours in a 3:1 H2:N2 gas mixture. The absorbent was then sealed in a vial containing a bed of 1 gram of NaOH and 1 gram of NH4Cl. The vial was kept at room temperature for 4 hours. The following reaction, shown in equation (4), produced NH3 gas absorbed into MnCl2, causing a discoloration from pink to dark brown. NaOH(s)+NH4Cl(s)=NaCl(s)+H2O(l)+NH3(g) (4)

[0309] Both 0.5 grams of catalyst and 1 gram of ammonium-treated absorbent, having a particle size of 150-400 μm, were homogeneously mixed and then added to the reaction chamber described in Example 12. In-situ pre-reduction at 350°C was carried out in the same manner as detailed in Example 12, but the stream of high-purity H2 was replaced with 80 sccm of a gas mixture of NH3:He:H2(vol). The presence of 5% NH3 was due to MnCl2·(NH3) during the pre-reduction of the catalyst. x The conditions were carefully maintained. The mixed catalyst-absorbent bed was then cooled to a reaction temperature of 300°C under the same 5:45:50 gas mixture of NH3:He:H2.

[0310] The results of pre-ammoniumization are shown in Figure 16. Figure 16 details the ammonia synthesis formation rate at 300°C and 20 barg in a mixed catalyst-absorbent bed for both ammonium and non-ammonium absorbents in an 80 sccm, 3:1 H2N2 flow. In general, pre-ammoniumization yielded 3.25 μmolg 触媒 -1 minutes -1The reaction exhibited high initial activity, reaching a high NH3 production rate, and a gradual decay to negligible activity over 20 hours on the stream (ToS). The same mixed bed configuration without pre-ammonium absorption and a stream of high-purity H2 during pre-reduction showed negligible activity from the start of the reaction initiation. Therefore, pre-ammonium mitigates the decomposition of metal halides during catalytic pre-reduction, resulting in considerable activity during the early ToS of NH3 synthesis, but still exhibits a limited lifetime as the reaction progresses.

[0311] The nature of this disclosure In some aspects, this disclosure relates to the following aspects:

[0312] A composition comprising multiple particles, each of which comprises a catalyst portion in close contact with an sorbent portion, thereby forming a catalyst-sorbent structure. A composition in which the sorbent portion is configured to remove ammonia when ammonia is essentially formed through a catalytic reaction by the catalyst portion, which is configured to convert nitrogen and hydrogen gases into a reaction mixture containing ammonia.

[0313] A composite material comprising a catalyst portion in close contact with an adsorbent portion, thereby forming a catalyst-adsorbent structure, wherein the adsorbent portion is configured to remove ammonia when ammonia is essentially formed through a catalytic reaction by the catalyst portion, which is configured to convert nitrogen and hydrogen gases into a reaction mixture containing ammonia.

[0314] A method for synthesizing ammonia from a gaseous feedstock containing hydrogen and nitrogen in the presence of a catalyst-sorbent structure comprising a catalyst portion in close contact with an sorbent portion, wherein the sorbent portion enables the removal of ammonia when ammonia is essentially formed via a catalytic reaction by the catalyst portion, which is configured to convert nitrogen and hydrogen gases into a reaction mixture containing ammonia.

[0315] An apparatus for producing ammonia from a feedstock containing hydrogen and nitrogen, wherein the apparatus includes a catalyst-sorbent structure comprising a catalyst portion in close contact with an sorbent portion, and the sorbent portion is configured to remove ammonia as ammonia is essentially formed through a catalytic reaction by the catalyst portion, which is configured to convert nitrogen and hydrogen gases into a reaction mixture containing ammonia.

[0316] A system for producing ammonia from a feedstock containing hydrogen and nitrogen, wherein the apparatus includes a catalyst-sorbent structure comprising a catalyst portion in close contact with an sorbent portion, the sorbent portion enabling the removal of ammonia as ammonia is essentially formed via a catalytic reaction by the catalyst portion, which is configured to convert nitrogen and hydrogen gases into a reaction mixture containing ammonia.

[0317] A catalyst-sorbent structure comprising a catalyst portion in close contact with an sorbent portion, wherein the sorbent portion is configured to remove ammonia when ammonia is essentially formed through a catalytic reaction by the catalyst portion, which is configured to convert nitrogen and hydrogen gases into a reaction mixture containing ammonia.

[0318] An integrated catalyst-sorbent structure comprising a catalyst portion and an sorbent portion, wherein the catalyst portion is capable of converting unreacted hydrogen feedstock and unreacted nitrogen feedstock into ammonia products, and the sorbent portion is capable of absorbing the generated ammonia, the conversion of the catalyst portion and the absorption of the sorbent portion can both occur at temperatures in the range of about 100°C to about 500°C, preferably about 200°C to about 400°C, more preferably about 250°C to about 380°C, and even more preferably about 280°C to about 350°C, and the conversion of the catalyst portion and the absorption of the sorbent portion can both occur at pressures in the range of about 2 bar to about 200 bar, preferably about 5 bar to about 100 bar, more preferably about 5 bar to about 50 bar, and even more preferably about 10 bar to about 40 bar.

[0319] Any of the above embodiments, wherein the catalyst and the sorbent portion are mixed together and pressed into a structural component such as pellets, tablets, granules, or extruded products.

[0320] Any of the above embodiments, wherein the adsorbent portion is first formed into a structural component such as a pressed pellet, tablet, granule, or extruded product to provide an absorbent core, and then a thin layer of active catalyst is coated as a surrounding shell or outer layer that at least encloses the absorbent core.

[0321] Any of the above embodiments, wherein the adsorbent portion is first formed into a structural component such as a pressed pellet, tablet, or extruded product to provide an absorbent core, a thin layer of active catalyst is coated as a surrounding shell or outer layer that at least encloses the absorbent core, and a second layer of absorbent is coated as a surrounding shell or outer layer that at least encloses the catalyst layer.

[0322] Any of the above embodiments, wherein the catalyst portion is first formed in a structural component such as a pressed pellet, tablet, granule, or extruded product to provide an active catalyst core, and then a layer of adsorbent is coated as a surrounding shell or outer layer that at least encloses the catalyst core.

[0323] Any of the above embodiments, wherein the sorbent portion and the catalyst portion are supported and dispersed sequentially or simultaneously along the same porous support, preferably by initial wetting impregnation, colloidal synthesis, or sol-gel method.

[0324] Any of the above embodiments, wherein the catalyst portion includes a discontinuous portion impregnated within the sorbent portion.

[0325] Any of the above embodiments, wherein the catalyst portion includes an active catalyst material comprising iron, cobalt, ruthenium, or a combination thereof.

[0326] Any of the above embodiments, wherein the sorbent portion comprises one or more metal halide absorbents having an absorption affinity for NH3 than for N2 and H2.

[0327] The catalyst-adsorbent structure further comprises a support material, preferably any of the above embodiments wherein at least a portion of the catalyst portion, at least a portion of the adsorbent portion, or a combination thereof, is supported on a molecularly porous support material.

[0328] Any of the above embodiments, wherein the catalyst-sorbent structure further comprises one or more promoters for increasing catalytic activity and / or improving catalytic stability.

[0329] Any of the above embodiments wherein the catalyst-sorbent structure has a porous structure having an average pore diameter of about 20 nm to about 50 microns, and in some preferred embodiments, about 50 nm to about 5 microns.

[0330] Catalyst-adsorbent mixture, approximately 1-1000 m 2Any of the above embodiments, wherein the catalyst-sorbent structure further comprises a molecular porous support material for the active catalyst so as to have a surface area in the range of / gram.

[0331] Any of the above embodiments wherein the molecular porous support material comprises an oxide selected from alumina, silica, magnesium, ceria, titania, or a combination thereof.

[0332] Any of the above embodiments wherein the active catalyst is porous and self-supporting.

[0333] Either of the above embodiments, wherein the sorbent portion supports (carries) the active catalyst, or the catalyst portion supports (carries) the sorbent portion.

[0334] Any of the above embodiments, wherein the adsorbent portion comprises one or more metal halide absorbents, the metal of the one or more metal halides is selected from Mn, Mg, Ca, Sr, and Fe, and the halide of the one or more metal halides is selected from Cl, Br, and I.

[0335] Any of the above embodiments, wherein the sorbent portion is a metal halide salt selected from the group consisting of LiCl, NH4Cl, CoCl2, MgCl2, CaCl2, MnCl2, FeCl2, NiCl2, CuCl2, ZnCl2, SrCl2, SnCl2, BaCl2, PbCl2, LiBr, NaBr, MgBr2, CaBr2, MnBr2, FeBr2, NiBr2, CoBr2, SrBr2, BaBr2, PbBr2, NaI, KI, CaI2, MnI2, FeI2, NiI2, SrI2, BaI2, and PbI2.

[0336] Any of the above embodiments, wherein the sorbent portion is a metal halide salt selected from the group consisting of MgCl2, CaCl2, MnCl2, and NiCl2.

[0337] Any of the above embodiments wherein the one or more metal halide salts include MnCl2, MgCl2, CaCl2, MgBr2, CaBr2, MgClBr, CaClBr, MgCaBr, and mixtures thereof.

[0338] Any of the above embodiments wherein the catalyst-sorbent structure is provided in the form of a structural component selected from pellets, tablets, granules, or extruded articles.

[0339] Any of the above embodiments wherein the adsorbent portion comprises one or more zeolites, preferably one or more aluminosilicate zeolites selected from zeolite Y, zeolite X, zeolite 4A, zeolite 5A, ZSM-5, or mixtures thereof.

[0340] The catalyst-sorbent structure comprises one or more coatings of the catalyst portion, the sorbent portion, or a combination thereof, and each coating has an average thickness of about 1 micron to about 200 microns, preferably about 10 microns to about 150 microns, more preferably about 20 microns to about 100 microns, in any of the above embodiments.

[0341] Any of the above embodiments, wherein the sorbent portion and the catalyst portion are supported and dispersed along the same porous support (sequentially or simultaneously, for example, by initial wetting impregnation, colloidal synthesis, or sol-gel method, or by other means) such that the catalyst-sorbent particles are supported on the same porous support.

[0342] Any of the above embodiments wherein the catalyst-sorbent structure has an average diameter of 1 to 20 mm, preferably 3 to 10 mm, and more preferably 3 to 9 mm.

[0343] Any of the above embodiments wherein the catalyst-sorbent structure has a partial sorbent load of 5% to 95% by weight, preferably 10% to 90% by weight, and more preferably 20% to 90% by weight.

[0344] Any of the above embodiments wherein the catalyst-sorbent structure has a catalyst partial loading amount of 0.01% to 20% by weight, preferably 0.25% to 10% by weight, and more preferably 0.5% to 5% by weight.

[0345] Any of the above embodiments wherein the catalyst-sorbent structure has a catalyst partial load of less than 5% by weight.

[0346] Any of the above embodiments wherein the catalyst-sorbent structure has a catalyst-sorbent weight ratio of approximately 1:1 to approximately 1:300, preferably approximately 1:1 to approximately 1:50, and more preferably approximately 1:1 to approximately 1:10.

[0347] The partial loading density of the adsorbent is approximately 100 kg / m³. 3 ~2000kg / m 3 The range is preferably about 300 kg / m 3 ~Approx. 1500kg / m 3 Within the range of approximately 500 kg / m², more comfortably. 3 ~Approx. 1200kg / m 3 Any of the aforementioned embodiments, which fall within the scope of the above.

[0348] The partial catalyst loading density is approximately 10 kg / m³. 3 ~2000kg / m 3 The range is preferably about 100 kg / m 3 ~Approx. 1500kg / m 3 In the range of approximately 150 kg / m², more preferably. 3 ~Approx. 1200kg / m 3 Any of the aforementioned embodiments, which fall within the scope of the above.

[0349] Any of the above embodiments wherein the catalyst-sorbent structure has a higher equilibrium NH3 conversion rate than that obtained when only the catalyst portion is used and not the sorbent portion, or the catalyst-sorbent structure has an improved product rate and / or reaction rate when only the catalyst portion is used and not the sorbent portion.

[0350] In any of the above embodiments, the catalyst-sorbent structure is provided in a reactor, preferably as a fixed bed such as a packed bed, and under normal operating conditions, the catalyst portion converts unreacted hydrogen and unreacted nitrogen into ammonia products, and the sorbent portion absorbs the generated ammonia.

[0351] Any of the above embodiments, wherein the catalyst-sorbent structure is placed in a reactor and the catalyst-sorbent structure is supported in a range of about 0.1% to about 99.9%, preferably about 10% to about 90%, and preferably about 15% to about 80%.

[0352] Any of the above embodiments, wherein the catalyst-sorbent structure is placed in a reactor, and the catalyst-sorbent structure is supported in the reactor in an amount of at least 10%, in some embodiments at least 20%, in some embodiments at least 30%, in some embodiments at least 40%, in some embodiments at least 50%, in some embodiments at least 60%, in some embodiments less than 90%, in some embodiments less than 80%, and in some embodiments less than 70% of the reactor volume.

[0353] Any of the above embodiments wherein the catalyst-sorbent structure is provided in the reactor in a catalyst portion having a weight range (w / w) of about 0.01% to about 20%, preferably about 0.25% to about 10%, more preferably about 0.5% to less than 5%.

[0354] Any of the above embodiments, wherein the catalyst-sorbent structure is provided in the reactor, and the catalyst portion is in the weight range of less than 5% by weight (w / w).

[0355] Any of the above embodiments wherein the process for producing ammonia using the integrated catalyst-sorbent structure has a process cycle in which the absorption capacity of the sorbent portion is less than the full absorption capacity. In some embodiments, the process cycle is at least 20% to about 95% of the full theoretical capacity, as defined by the operating temperature and operating pressure of the process bed.

[0356] A process for producing ammonia using the integrated catalyst-sorbent structure comprises an initial process cycle having an initial conversion rate and a second process cycle having a second conversion rate, wherein the second conversion rate is lower than the initial conversion rate, in some embodiments at least 1% lower than the initial conversion rate, and in some preferred embodiments 1% to 10% lower than the initial conversion rate, in any of the above embodiments.

[0357] Any of the above embodiments, wherein the process for generating ammonia comprises providing the integrated catalyst-sorbent structure to a plurality of beds, the plurality of beds being provided in series, parallel, or a combination of series and parallel.

[0358] Any of the above embodiments wherein the unreacted hydrogen is supplied from a hydrogen source including an electrolytic cell.

[0359] Any of the above embodiments wherein the unreacted nitrogen is supplied from a nitrogen source, and the nitrogen source is a pressure swing adsorption (PSA) system, an air separation unit (ASU) system, a membrane separator, a nitrogen tank, or a combination thereof.

[0360] The one or more zeolites have a binding affinity for NH3 than for N2 and H2, preferably at least 5 times, in some embodiments at least 10 times, in some embodiments at least 100 times, in some embodiments at least 200 times, in some embodiments at least 300 times, in some embodiments at least 400 times, in some embodiments at least 500 times, in some embodiments at least 600 times, in some embodiments at least 700 times, and in some embodiments at least 1000 times or more, having an affinity for NH3 than for N2 and / or H2, in any of the aforementioned embodiments.

[0361] Any of the above embodiments wherein one or more zeolites have pore sizes smaller than N2 and H2 molecules but larger than NH3 molecules, and as a result, the pore size results in size exclusion of unreacted N2 and unreacted H2 but allows for the flow of NH3.

[0362] Any of the above embodiments, wherein one or more zeolites have a pore size of about 3 Å to about 5 Å, preferably about 4 Å to about 5 Å.

[0363] Any of the above embodiments, wherein one or more zeolites have a pore size larger than that of an NH3 molecule.

[0364] Any of the above embodiments, wherein one or more zeolites have a pore size greater than 5 Å.

[0365] Any of the above embodiments, wherein the one or more zeolites have a desired pore size formed by ion exchange for partially or completely replacing Na and / or H cations with one or more other metals, preferably one or more alkali metals or transition metals.

[0366] Any of the above embodiments, wherein the adsorbent portion comprises one or more zeolites, preferably one or more aluminosilicate zeolites, and the one or more zeolites are supported on the catalyst portion.

[0367] Any of the above embodiments further comprising a secondary sorbent portion.

[0368] Any of the above embodiments, wherein the secondary sorbent portion contains one or more metal halides.

[0369] Any of the above embodiments, wherein the secondary sorbent portion includes one or more zeolites different from the sorbent portion.

[0370] Any of the above embodiments further comprising a promoter material, wherein the promoter material preferably comprises K, Ce, Cs, Ba, or a mixture thereof.

[0371] The promoter material is preferably supported in one or more zeolites in an amount of more than 0 to a maximum of about 10% by weight, in any of the above embodiments.

[0372] Any of the above embodiments wherein the integrated catalyst-sorbent structure is in the form of pellets, tablets, extruded articles, or granules having an average diameter of 1 mm to 20 mm, preferably 3 mm to 10 mm, and more preferably 3 mm to 9 mm.

[0373] Any of the above embodiments, wherein the integrated catalyst-adsorbent structure is in the form of a monolithic structure having a support material, and the support material is preferably a ceramic material, a metal oxide such as alumina, or a combination thereof.

[0374] Various embodiments of systems, devices, and methods are described herein. These embodiments are given only as examples and are not intended to limit the scope of the claimed invention. Furthermore, it should be understood that various features of the described embodiments can be combined in various ways to produce a number of additional embodiments. In addition, various materials, dimensions, shapes, configurations, and positions, etc., are described for use with the disclosed embodiments, but anything other than those disclosed can be used without exceeding the scope of the claimed invention.

[0375] Those skilled in the art will recognize that the subject matter of this specification may include fewer features than those illustrated in the individual embodiments described above. The embodiments described herein are not intended to be an exhaustive presentation of how various features of the subject matter of this specification can be combined. Thus, the embodiments are not mutually exclusive combinations of features, but rather, as will be understood by those skilled in the art, various embodiments may include different combinations of individual features selected from different individual embodiments. Furthermore, elements described in relation to one embodiment may be implemented in other embodiments even if they are not described in that embodiment, unless otherwise specified.

[0376] A dependent claim may refer to a specific combination of one or more other claims in a claim, and other embodiments may also include a combination of a dependent claim with the subject matter of each other dependent claim, or a combination of one or more features with other dependent or independent claims. Such combinations are proposed herein unless otherwise stated that a specific combination is not intended.

[0377] Any incorporation by reference of the above documents is limited so as not to include any subject matter contrary to the express disclosure herein. Any incorporation by reference of the above documents is further limited so as not to include the claims in the documents incorporated herein by reference. Any incorporation by reference of the above documents is further limited so as not to include any definitions provided in the documents unless expressly included herein by reference.

[0378] For the purpose of interpreting the claims, it is expressly intended that the provisions of Section 112(f) of the U.S. Patent Act should not be invoked unless the specific terms “means for” or “steps for” are explicitly stated in the claims.

Claims

1. Catalyst-adsorbent structure, The catalyst portion is in direct contact with the sorbent portion, the catalyst portion comprises one or more active catalysts, and the sorbent portion comprises one or more sorbents. The catalyst portion is capable of converting unreacted hydrogen and unreacted nitrogen into ammonia products via a catalytic reaction. Essentially, a catalyst-adsorbent structure in which, when the ammonia product is formed via the catalytic reaction, the adsorbent portion allows the ammonia product to be removed from the catalyst portion to the adsorbent portion.

2. The catalyst-adsorbent structure according to claim 1, wherein the catalyst portion and the adsorbent portion are in the form of pressed pellets, tablets, extruded articles, or monolithic structures.

3. The catalyst-sorbent structure according to claim 1, further comprising a porous support material selected from alumina, silica, magnesium, ceria, titania, or a combination thereof, wherein the sorbent portion and the catalyst portion are supported and dispersed on the porous support material.

4. The catalyst-adsorbent structure according to claim 1, wherein the one or more active catalysts include iron, cobalt, ruthenium, molybdenum, or a combination thereof.

5. The catalyst-adsorbent structure according to claim 4, wherein the one or more adsorbents comprises one or more metal halide absorbents, the metal of the one or more metal halides comprises Mn, Mg, Ca, Sr, Fe, or a mixture thereof, and the halide of the one or more metal halides comprises Cl, Br, I, or a mixture thereof.

6. The one or more sorbents comprises one or more metal halide absorbents, and the one or more metal halide absorbents comprises MnCl ２ MgCl ２ CaCl ２ MgBr ２ CaBr ２ The catalyst-adsorbent structure according to claim 4, comprising MgClBr, CaClBr, MgCaBr, or a mixture thereof.

7. The catalyst-sorbent structure according to claim 4, wherein the one or more sorbents comprise one or more zeolites.

8. The catalyst-adsorbent structure according to claim 7, wherein the one or more zeolites are selected from zeolite Y, zeolite X, zeolite 4A, zeolite 5A, ZSM-5, or a mixture thereof.

9. The catalyst-sorbent structure according to claim 7, wherein one or more zeolites have a binding affinity for the ammonia product, and the binding affinity is at least 10 times greater than that for the unreacted hydrogen supply material and the unreacted nitrogen supply material.

10. The catalyst-sorbent structure according to claim 7, wherein at least a portion of one or more zeolites has a pore size smaller than the unreacted hydrogen supply material and the unreacted nitrogen supply material but larger than the ammonia product, and as a result the pore size is configured to achieve size exclusion of the unreacted hydrogen supply material and the unreacted nitrogen supply material, while allowing the flow of the ammonia product.

11. The catalyst-adsorbent structure according to claim 10, wherein at least a portion of the one or more zeolites has a desired pore size formed by ion exchange that partially or completely replaces at least a portion of the one or more cations.

12. The catalyst-sorbent structure according to claim 7, wherein one or more zeolites have a pore size smaller than that of the ammonia product.

13. The catalyst-adsorbent structure according to claim 7, further comprising a secondary adsorbent portion containing one or more metal halide absorbents, one or more zeolites, or a combination thereof.

14. The catalyst-sorbent structure according to claim 7, further comprising one or more accelerator materials, wherein the one or more accelerator materials comprise K, Ce, Cs, or a mixture thereof.

15. The catalyst-sorbent structure according to claim 1, wherein the catalyst-sorbent structure includes a porous structure having an average pore diameter of about 20 nm to about 5 microns.

16. A method for synthesizing ammonia from a gaseous feedstock, A step of providing a catalyst-sorbent structure in a reactor bed, wherein the catalyst-sorbent structure includes a catalyst portion in direct contact with an sorbent portion, the catalyst portion includes one or more active catalysts, and the sorbent portion includes one or more sorbents. A step of providing the catalyst-sorbent structure to a gaseous supply material including an unreacted hydrogen supply material and an unreacted nitrogen supply material to cause a catalytic reaction near the catalyst portion, wherein the catalytic reaction converts a portion of the unreacted hydrogen supply material and a portion of the unreacted nitrogen supply material into ammonia products, A method wherein, essentially, when the ammonia product is formed via the catalytic reaction, the sorbent portion allows the ammonia product to be removed from the catalytic portion to the sorbent portion.

17. The method according to claim 16, further comprising desorbing the ammonia product from the sorbent portion.

18. The method according to claim 16, wherein the catalytic reaction that forms the ammonia product and the removal of the ammonia product to the sorbent portion can be carried out at a temperature in the range of about 280°C to about 400°C and a pressure in the range of about 5 bar to about 50 bar.

19. The method according to claim 16, wherein the one or more sorbents comprises one or more metal halide absorbents or one or more zeolites.

20. The method according to claim 19, wherein the one or more active catalysts include iron, cobalt, ruthenium, molybdenum, or a combination thereof.

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