Process for gas phase allothermic catalytic conversions and reactor therefor

Abstract

A process for a gas phase allothermic catalytic conversion comprises providing a reactor having a first chamber, a reaction zone in fluid communication with the first chamber and a second chamber in fluid communication with the reaction zone, the reaction zone containing a catalyst material and being electrically heated, each of the first and second chambers containing a solid heat exchange material and being selectively operable in a cooling and a heating mode. A feed is introduced into a selected one of the first and second chambers when the selected chamber is in heating mode and the other of the first and second chamber is in cooling mode. The feed gas is conducted through the reaction zone so that the feed gas contacts the catalyst material and forms a gaseous reaction product. The gaseous reaction product is conducted through the other of the first and second chamber. The direction of gas flow through the reactor is periodically reversed so that the first chamber and the second chamber alternately operate in the cooling and heating modes. Further provided is an apparatus for a gas phase allothermic catalytic conversion, suitable for carrying out the process. The process and apparatus allow for energetically efficient allothermic conversion processes with reduced reliance on reactor feed / effluent heat exchange.

Description

[0001] Process for Gas Phase Allothermic Catalytic Conversions and Reactor Therefor

[0002] The present invention relates to a process for a gas phase allothermic catalytic conversion, in particular catalytic reforming of ammonia to recover hydrogen, and a reactor therefor.

[0003] In today's chemical industry, heat is supplied in allothermic conversion processes, such as steam reforming or steam cracking, by means of the combustion of fossil fuels, for example by burning natural gas. Combustion of fossil fuels is associated with high emissions of CO2 and, hence, these technologies will be subject of increasing pressure from authorities and lawmakers in the coming decades. An alternative to fossil fuels is to provide the heat required for allothermic conversion processes by electric energy, especially coming from renewable sources having a low CO2 footprint. Since in most countries electricity is currently a more expensive energy source than fossil fuels, and in order to increase the overall efficiency of such an alternative electrically heated allothermal conversion process, a high efficiency of the utilized electricity is preferred.

[0004] Fired box processes are almost invariably associated with heat loss through the effluent heat of the off gas. Though back-integration of the associated heat streams, including the effluent heat in the off gas of fired box processes, heat integration may be possible to a certain extent. However, the resulting processes become complex, inflexible, necessarily comprises multiple heat exchangers and require high capital expenditures. In contrast, electrical heating allows for a more efficient energy utilization in a chemical process.

[0005] An example of an allothermic process is the decomposition, or reforming, of ammonia to produce nitrogen and hydrogen. In the decomposition of ammonia to nitrogen and hydrogen, the sensible heat content of the dissociated gas, at any temperature level, is greater than that of the corresponding quantity of ammonia gas from which it was formed. Hence, even with a perfect heat exchanger, it is not possible to produce sufficient exit gas cooling merely by exchange against incoming ammonia feed.

[0006] WO 2023 / 174935 A1 relates to a method for the production of nitric oxide from a gaseous reactant mixture containing oxygen and nitrogen in a reactor comprising a reaction zone with a heat input device and at least two regenerator zones, each regenerator zone having a low temperature section on one end and a high temperature section at the other end of the regenerator zone, the high temperature sections being fluidically connected to the reaction zone. The method comprises supplying heat through the heat input device to the reaction zone until a temperature of 1500°C to 2500°C is reached in the reaction zone; passing the reactant mixture through a first regenerator zone into the reaction zone in which the reactant mixture reacts to form a product mixture, passing the product mixture from the reaction zone through a second regenerator zone and withdrawing at least part of the product mixture from the second regenerator zone; reversing the direction of flow and passing the reactant mixture through the second regenerator zone into the reaction zone in which the reactant mixture reacts to form a product mixture, passing the product mixture from the reaction zone through the first regenerator zone and withdrawing at least part of the product mixture from the first regenerator zone; and reversing the direction of flow and periodically. The high temperature sections of the regenerator zones comprise a plurality of channels with a hydraulic diameter of 0.5 mm to 5 mm each, the inner walls of which are made of oxide ceramics.

[0007] Eigenberger, Gerhart, and Ulrich Nieken. "Catalytic combustion with periodic flow reversal." Tenth International Symposium on Chemical Reaction Engineering. Pergamon, 1988 describe the design features of a reverse flow reactor and operating conditions that are suitable for the targeted adjustment of the operating state, in particular the manipulation of the axial temperature profile. A suitable measure for controlling the outlet temperature from the reactor is hot gas extraction from the center of the reaction zone. A small proportion of the reaction gas is extracted from the hot area of the reactor. As a result, the heat generated in the reactor can be almost completely extracted at a high temperature level and used to generate energy. In addition to energy utilization, this measure has another advantage: the reaction medium exits the reactor outlet at low temperatures and is therefore easier to handle.

[0008] US 7,811 ,529 discloses a process for obtaining hydrogen from ammonia in a membrane reactor, where in a first step the ammonia is evaporated and in a second step it is reformed in a hydrogen membrane reactor, with the resultant hydrogen being removed at the same time by means of a membrane. Ammonia and the gaseous retentate from the membrane process are burned in a burner with air, thereby providing the necessary heat for the evaporation and reforming via heat exchange.

[0009] WO 2023 / 016879 relates to a method for obtaining hydrogen from methanol or ammonia, e.g., for operating fuel cells. Methanol or ammonia is evaporated and is reformed into a hydrogen-containing gas mixture. Subsequently, hydrogen is separated from the gas mixture in a membrane process, and the gaseous retentate of the membrane process is combusted using ambient air. The combustion gases are guided via at least two different heat exchangers in order to provide (I) the reaction heat for reforming the methanol or ammonia and then (ii) the evaporation heat for evaporating the reformer feed, and the permeate of the membrane process preheats the ambient air for the burner in a heat exchanger.

[0010] There remains a need for energetically efficient allothermic conversion processes with reduced reliance on reactor feed I effluent heat exchange.

[0011] Summary of the Invention

[0012] The present invention provides a process and an apparatus for a highly efficient allothermic catalytic conversion process, using an electrically heated reverse flow reactor. By periodical reversal of the flow direction a high level of heat integration can be achieved in the reverse flow reactor itself, making optimal use of the supplied electrical energy. Furthermore, by extracting a fraction of the gaseous reaction product via a side draw the temperature profiles inside the electrically heated reverse flow reactor can be further optimized and excess heat can be utilized at high temperature levels. Utilization of the electrically heated reverse flow reactor concept, including the integration in a detailed production process, is demonstrated for the catalytic reforming of ammonia to produce hydrogen.

[0013] The invention relates to a process for a gas phase allothermic catalytic conversion, comprising a) providing a reactor having a first chamber, a reaction zone in fluid communication with the first chamber and a second chamber in fluid communication with the reaction zone, the reaction zone containing a catalyst material and being electrically heated, each of the first and second chambers containing a solid heat exchange material and being selectively operable in a cooling and a heating mode; b) introducing a feed gas into a selected one of the first and second chambers when the selected chamber is in heating mode and the other of the first and second chamber is in cooling mode; c) conducting the feed gas through the reaction zone so that the feed gas contacts the catalyst material and forms a gaseous reaction product; d) conducting the gaseous reaction product through the other of the first and second chamber; e) periodically reversing the direction of gas flow through the reactor so that the first chamber and the second chamber alternately operate in the cooling and heating modes.

[0014] A chamber operated in "heating mode" denotes the chamber that contains preheated solid heat exchange material, wherein the feed gas is heated and the heat exchange material is cooled. A chamber operated in "cooling mode" denotes the chamber that contains unheated solid heat exchange material, wherein the hot gaseous reaction product is cooled and heat exchange material is heated.

[0015] In an embodiment, the process comprises discharging a portion of the gaseous reaction product from the reaction zone or from the one of the first or second chamber operating in cooling mode at an intermediate point between the reaction zone and the distal end of the chamber. For example, 5 to 30 vol. -%, preferably 10 to 20 vol.-%, of the gaseous reaction product may thus be discharged.

[0016] By discharging a portion of the gaseous reaction product sufficient heat may be withdrawn to maintain the temperature of the gaseous reaction product leaving the chamber via the distal end of the chamber below a predetermined temperature.

[0017] In one embodiment, the process comprises, prior to reversing the direction of gas flow, stopping the flow of feed gas into the selected one of the first and second chambers, and introducing a purge gas into the selected chamber, thereby purging the selected chamber and pushing feed gas which has remained in the selected chamber through the reaction zone. In a particular aspect of the above embodiment, the process comprises providing a supply of the feed gas and conducting the feed gas through a buffer tank; wherein when the flow of feed gas into the selected one of the first and second chambers is stopped, feed gas accumulates in the buffer tank, thereby increasing pressure in the buffer tank, and when the flow of feed gas into the other of the first and second chamber is resumed, pressure is relieved in the buffer tank.

[0018] The present invention also provides, in another aspect thereof, an apparatus for carrying out a process as defined above.

[0019] The apparatus of the invention includes a reactor having a first chamber, a reaction zone in fluid communication with the first chamber and a second chamber in fluid communication with the reaction zone, the reaction zone containing a catalyst material and being electrically heated, each of the chambers containing a solid heat exchange material and being selectively operable in a cooling and heating mode; the reactor being adapted for introducing a feed gas into a selected one of the first and second chambers when the selected chamber is in heating mode and the other of the first and second chamber is in cooling mode; conducting the feed gas through the reaction zone so that the feed gas contacts the catalyst material and forms a gaseous reaction product; and conducting the gaseous reaction product through the other of the first and second chamber; the apparatus further including gas flow directing means for periodically reversing the direction of gas flow through the reactor so that the first chamber and second chamber alternately operate in the cooling and heating modes.

[0020] The apparatus may comprise outlet means in fluid communication with at least one of

[0021] - the reaction zone for discharging a portion of the gaseous reaction product from the reaction zone; and

[0022] - each of the first and second chamber for discharging a portion of the gaseous reaction product from the one of the first or second chamber operating in cooling mode at an intermediate point between the reaction zone and the distal end of the chamber.

[0023] In one embodiment, the reactor is further adapted for temporarily stopping the feed gas stream and introducing a purge gas stream into the selected one of the chambers before the direction of gas flow through the reactor is reversed.

[0024] In one embodiment, the apparatus further includes a buffer tank for temporarily storing a feed gas and a reactor; the buffer tank being adapted to release the feed gas to the reactor. Detailed Description

[0025] In the process of the invention, the feed gas is passed through a selected one of the first and second chambers. The selected chamber contains preheated solid heat exchange material wherein the feed gas is heated to a desired temperature. Subsequently, the feed gas is passed through an electrically heated reaction zone containing a catalyst material, and finally through the other of the first and second chamber containing an unheated solid heat exchange material which is similar to that in the selected chamber. As the gases continue to flow in one direction, the temperature of the heat exchange material in the selected chamber drops as it gives up its heat to the incoming feed gas. Correspondingly, the hot gaseous reaction product being cooled by the unheated heat exchange material in the other chamber heats up this heat exchange material. The heat exchange material of the selected chamber soon loses its ability to heat the incoming feed gas sufficiently, and the heat exchange material of the other chamber loses its ability to cool the gaseous reaction product sufficiently. Preferably, before this point is reached, the flow of gas is reversed in direction so that the relatively cool feed gas is heated by the hot heat exchange material previously heated by the hot gaseous reaction product, whereas the hot gaseous reaction product is cooled by the heat exchange material which had previously given up its heat to the incoming feed gas. This reversal of gas flow is made to occur at intervals so as to maintain essentially optimum conditions.

[0026] The flow of gas may be reversed in direction by means of flow directing means such as a four-way switch valve or other valving means.

[0027] In an embodiment, the process comprises discharging a portion of the gaseous reaction product from the reaction zone. In an embodiment, the process comprises discharging a portion of the gaseous reaction product from the one of the first or second chamber operating in cooling mode at an intermediate point between the reaction zone and the distal end of the chamber (also referred to as "side draw” in the following). The "distal end” denotes the end of the first or second chamber opposite the reaction zone through which the feed gas enters the selected chamber when operated in heating mode or through which (the majority of) the gaseous reaction product leaves the chamber when operated in cooling mode. By discharging a portion of the hot gaseous reaction product sufficient heat may be withdrawn so as to maintain the reaction products at the outlet below a predetermined maximum temperature. As the heat capacities of the feed gas and the gaseous reaction product may not match exactly, a temperature build-up may occur and the temperature build-up can exceed that desired for optimum operation. Hence, a discharge may be provided so as to permit a portion of the hot gaseous reaction product to leave the reaction zone or preferably the first or second chamber at an intermediate point between the reaction zone and the distal end of the chamber. The portion of the gaseous reaction product may be withdrawn at a point adjacent to the reaction zone, just after the reaction has taken place. The portion of the gaseous reaction product may be withdrawn at a point farther away from the reaction zone. For example, the point of withdrawal is located within 0 to 80%, preferably 0 to 60%, more preferably 0 to 40%, of the length of the chamber, seen from the reaction zone. The discharge being used will vary with the direction of gas flow, always being the one which will remove the hot gaseous reaction product just after the reaction. The amount removed at this point can be varied to maintain the temperature of the gaseous reaction product leaving the chamber operating in cooling mode via the distal end of the chamber below a predetermined temperature.

[0028] The heat of the discharged portion of the gaseous reaction product is recovered at a high temperature stage. It can be recovered by heat exchange and used elsewhere in the process, for example for steam generation, such as steam generation for an integrated production facility, for downstream processes such as distillation units, heating of other chemical reactions, ammonia preheating or evaporation.

[0029] In one embodiment, the process comprises, prior to reversing the direction of gas flow, stopping the flow of feed gas into the selected one of the first and second chambers, and introducing a purge gas into the selected chamber, thereby purging the selected chamber and pushing feed gas which has remained in the selected chamber through the reaction zone.

[0030] Thus, the process may comprise purge periods before the flow direction of the feed gas is reversed. A purge period involves stopping the flow of feed gas into the selected one of the first and second chambers, and introducing a purge gas into the selected chamber, thereby purging the selected chamber and pushing feed gas which has remained in the selected chamber through the reaction zone, thereby forming an additional quantity of gaseous reaction product. Purging serves to avoid contamination of the gaseous reaction product with unreacted feed gas after the flow direction has been reversed.

[0031] In one embodiment, the process comprises sensing the flow rate of feed gas upstream of the reactor and adjusting delivery of feed gas to the reactor, e.g., via a throttle member arranged in the duct of the flow of feed gas to the reactor, according to the sensed flow rate.

[0032] In one embodiment, the process comprises sensing the flow rate of purge gas upstream of the reactor and adjusting delivery of purge gas to the reactor, e.g., via a throttle member arranged in the duct of the flow of purge gas to the reactor, according to the sensed flow rate.

[0033] The purge gas may be selected from an inert gas, a partial flow of the gaseous reaction product, a partial flow of a process gas from a downstream processing stage, and a regeneration gas. Suitable inert gases include nitrogen, argon, water vapor and carbon dioxide. In the case of reforming of ammonia, for example, a recirculated partial flow of the gaseous reaction product may be a gas mixture containing hydrogen, nitrogen and ammonia. In the case of steam reforming of natural gas, for example, a recirculated partial flow of the gaseous reaction product can be a gas mixture containing methane, water vapor, carbon monoxide and carbon dioxide. The partial flow of a process gas returned from its downstream processing stage can, for example, be a gas containing hydrogen or pure nitrogen in the case of reforming of ammonia, or a gas containing hydrogen or a gas containing hydrogen and carbon monoxide in the case of reforming of ammonia. The regeneration gas may be a gas containing oxygen, water vapor and / or carbon dioxide.

[0034] Interruption of the productive periods of the reactor by purge periods results in the reactor using feed gas in short bursts. In particular, the reactor alternates between two phases. During an active phase, the reactor actively consumes feed gas. The purge period is an inactive phase, during which the reactor consumes only a limited amount of or essentially no feed gas. A "limited amount of feed gas" is understood to constitute less than 50 vol.-% of the amount of feed gas consumed in the active phase, for example less than 30 vol.-%, such as less than 10 vol.-%. In a particular embodiment, the consumer essentially does not consume feed gas during the inactive phase, for example less than 1 vol.-%, or 0 vol.-% of the amount of feed gas consumed in the active phase. This leads to a situation where the supply must respond rapidly to changes in demand. Often, however, there is a mismatch between the sudden demand for feed gas by a reactor and the supply system's ability to respond to that demand.

[0035] For example, the feed gas supply may be provided by vaporization of a liquid raw material. An evaporator may be controlled only with large time constants, meaning that it can ramp up or down only slowly. This is due to the physical characteristics and operational limitations of the evaporator.

[0036] Evaporators work best if the liquid is evaporated against an essentially constant pressure. Variable pressure operation can induce stress on the evaporator components such as the expansion valve and compressor due to rapid pressure changes, which could lead to increased wear and potential mechanical failure over time. Rapid changes in demand of vaporized feed gas may however cause pressure fluctuations.

[0037] To accommodate intermittent demand and slow supply response, the system may require buffer solutions. Such buffer solutions may include storage capabilities that allow feed gas to be stored during low demand times for use during peak demand.

[0038] In a particular aspect of the above embodiment, the process comprises providing a supply of the feed gas and conducting the feed gas through a buffer tank; wherein when the flow of feed gas into the selected one of the first and second chambers is stopped, feed gas accumulates in the buffer tank, thereby increasing pressure in the buffer tank, and when the flow of feed gas into the other of the first and second chamber is resumed, pressure is relieved in the buffer tank.

[0039] In a preferred embodiment, the process comprises vaporizing a liquid feed to provide the supply of the feed gas, in particular in an evaporator. In this particular embodiment, the feed gas constitutes a feed vapor. The evaporator can be any suitable evaporator. Examples of suitable evaporators include tubular evaporators (horizontal and vertical) and plate-type evaporators, such as forced circulation tubular evaporators, falling film evaporators including falling film tubular evaporators (horizontal and vertical), rising / falling film evaporators including rising / falling film tubular evaporators, gasketed plate evaporators, forced circulation evaporators including forced circulation tubular evaporators and forced circulation plate evaporators, and agitated thin-film evaporators such as wiped film evaporators.

[0040] In one embodiment, the process comprises sensing the flow rate of liquid to the evaporator and adjusting delivery of liquid to the evaporator, e.g., via a pump, according to the sensed flow rate.

[0041] Upstream of the evaporator, a preheater may be provided for obtaining a preheated liquid, and a vaporized feed gas is obtained from the preheated liquid in the evaporator. Preheating the liquid reduces the amount of energy required by the evaporator to reach the necessary boiling point. Since part of the heating required to vaporize the liquid takes place in the preheater, the evaporator can operate more efficiently. In particular, by raising the temperature of the liquid before it enters the evaporator, the system can achieve higher vaporization rates, allowing for more efficient processing and potentially increasing throughput. If waste heat from other processes or operations is available, it can be effectively used for preheating, allowing to maximize the energy efficiency of the system.

[0042] In one embodiment, the process involves sensing the temperature of preheated liquid upstream of the evaporator and controlling the energy input in the preheater based on the temperature sensed in the preheated liquid.

[0043] The buffer tank for temporarily storing the feed gas serves a critical role in this process. Thus, during the inactive phases of the reactor, feed gas accumulates in the buffer tank. As feed gas is supplied and not consumed, pressure within the buffer tank increases, creating a reservoir of feed gas that can be accessed as soon as the reactor resumes its active phase.

[0044] During the active phase, the pressure in the buffer tank is relieved to supply the reactor with the required feed gas. The design of the pressure relief mechanism is essential as it allows gas to flow to the reactor without abrupt pressure changes, which could otherwise lead to operational issues such as flow disruptions or equipment damage. It may involve valves, pressure regulators, or automated systems that respond to the dynamic needs of the reactor. In one embodiment, the process comprises sensing the feed gas pressure upstream the buffer tank and adjusting delivery of feed gas to the buffer tank, e.g., via a throttle member arranged in the duct of the flow of feed gas to the buffer tank, according to the sensed pressure.

[0045] This buffer system provides operational benefits. By smoothing out the supply of feed gas and aligning it with the reactor's demand cycles, the process can enhance overall system efficiency. By preventing rapid fluctuations in pressure and flow, the system reduces mechanical stress on pipelines and valves, thereby prolonging their lifespan. The volume of the buffer tank depends on the feed gas demand of the reactor and the envisaged reaction underlying the process. Generally, the volume of the buffer tank is in the range of 0.1 to 500 m3, such as 0.2 to 400 m3, for example 0.3 to 300 m3or 0.5 to 200 m3.

[0046] The temperature in the buffer tank depends on the composition of the feed gas. Generally, the temperature in the buffer tank is in the range of -30 to 500 °C, such as 0 to 400 °C, for example 20 to 300 °C.

[0047] In the process of the invention, the periodic feed gas accumulation and release of feed gas from the buffer tank does not necessarily require discrete pumps arranged between the feed gas supplier, e.g., an evaporator and the buffer tank, or between the buffer tank and the reactor. Rather, conducting the feed gas to and through the buffer tank, and from the buffer tank to the reactor may be achieved by a sufficient overpressure of the feed gas supply line upstream of the buffer tank relative to the feed gas flow line from the buffer tank to the reactor.

[0048] In one embodiment, the pressure difference between the feed gas supply line upstream of the buffer tank and the feed gas flow line from the buffer tank to the reactor is in the range of 0.05 to 50 bar, preferably 0.1 to 30 bar, calculated as the pressure in the feed gas supply line upstream of the buffer tank minus the pressure in the feed gas flow line from the buffer tank to the reactor.

[0049] Likewise, accumulation of feed gas in the buffer tank involves a sufficient overpressure of the feed gas supply line upstream of the buffer tank relative to the buffer tank, and relief of pressure in the buffer tank to the reactor requires a sufficient overpressure of the buffer tank relative to the feed gas flow line to the reactor.

[0050] In one embodiment, the pressure difference between the feed gas supply line upstream of the buffer tank and the buffer tank is in the range of 0.05 to 50 bar, preferably 0.1 to 30 bar, calculated as the pressure in the feed gas supply line upstream of the buffer tank minus the pressure in the buffer tank.

[0051] In one embodiment, the pressure difference between the buffer tank and the feed gas flow line to the reactor is in the range of 0.05 to 50 bar, preferably 0.1 to 30 bar, calculated as the pressure in the buffer tank minus the feed gas flow line to the reactor.

[0052] The pressure in the feed gas supply line upstream of the buffer tank is not especially limited and may be in the range of 0.05 to 140 bar, such as 0.1 to 90 bar, for example 0.2 to 60 bar.

[0053] The pressure in the feed gas flow line from the buffer tank to the reactor is not especially limited and may be in the range of 0.1 to 150 bar, such as 0.2 to 100 bar, for example 0.3 to 70 bar. Periodically increasing and relieving the pressure induces a pressure oscillation in the buffer tank. The amplitude of the pressure oscillation is preferably in the range of 0.05 to 50 bar, more preferably 0.1 to 30 bar.

[0054] Typically, the supply of feed gas is controllable with a time constant larger than the time constant for controlling the flow of feed gas from the buffer tank to the reactor.

[0055] Most physical situations involve exponential growth or decay where the rate of change of a variable is proportional to its current value. The transition of the system is asymptotic, and follows or can be approximated by an exponential function: wherein t is time; y(t) is a variable that describes the system status (status variable); y(0) is the initial value of the state variable before the disturbance; y(°°) is the value of the state variable in the new steady state; and T is the time constant of the system.

[0056] The time constant (T) is defined as the amount of time required for an initial variable (e.g., pressure or volume flow) to decay or rise by 63.2% in route to steady state at the new variable. y(r) - y(0)

[0057] , - — = (1 - e1) « 0.632 y(o°) — y(0)

[0058] Typically, the time constant for the supply of feed gas, e.g., the start-up and shutdown of an evaporator, is in the range of 30 to 10,000 s, such as 60 to 5,000 s.

[0059] In one embodiment, the process comprises sensing the flow of feed gas from the buffer tank to the reactor and controlling the flow according to the sensed flow, preferably to maintain an essentially constant flow of feed gas from the buffer tank to the reactor during the active phase. The time constant for sensing the flow of feed gas from the buffer tank to the reactor and controlling the flow according to the sensed flow is preferably in the range of 0.1 to 10 s, more preferably 0.2 to 7 s, even more preferably 0.3 to 5 s and most preferably 0.5 to 3 s.

[0060] Incorporating a sensing and actuator mechanism helps manage the flow of feed gas from the buffer tank to the reactor. Controlling the flow of feed gas based on the sensed data may involve various methodologies. In one embodiment, the gas flow control can be automated through actuated valves that respond to signals from the sensing apparatus. For example, if the system detects that the flow is lower than desired, it can adjust the valve to increase the flow, ensuring the reactor receives the necessary amount of feed gas. The primary goal is to maintain an essentially constant flow of feed gas during the active phase. Many industrial processes, such as combustion or chemical reactions, require a stable flow of feed gas to ensure optimal reaction conditions and efficiency. By maintaining a constant flow, the system minimizes the risk of surges that can impact equipment or lead to inefficiencies. Conversely, dips in flow can disrupt operations, which can be costly and potentially hazardous.

[0061] In one embodiment, the process comprises sensing the feed gas pressure upstream of the buffer tank and adjusting the delivery of feed gas to the buffer tank, e.g., via a throttle member arranged in the duct of the flow of feed gas to the buffer tank, according to the sensed pressure, preferably to maintain an essentially constant pressure upstream of the buffer tank. The time constant for sensing the feed gas pressure upstream of the buffer tank and controlling the flow of feed gas to the buffer tank according to the sensed pressure is preferably in the range of 0.1 to 10 s, more preferably 0.2 to 7 s, even more preferably 0.3 to 5 s and most preferably 0.5 to 3 s.

[0062] In this embodiment, the system includes a pressure sensor positioned upstream of the buffer tank. The sensed pressure becomes an input for the control mechanisms of the gas delivery system. This feedback loop allows the system to adjust gas flow to the buffer tank in response to pressure changes, which aids in maintaining optimal conditions for system operation. The pressure management mitigates the risk of overpressurization in the supply lines upstream of the buffer tank.

[0063] The goal of maintaining an essentially constant pressure upstream of the buffer tank may be important for several reasons. For example, when the feed gas is obtained by vaporizing a fluid, the significance of maintaining constant pressure becomes pronounced, as vaporization processes often rely on specific pressure conditions for optimal performance. Evaporators work most effectively when they are fed at a stable pressure. Fluctuations in pressure can lead to inconsistent vaporization rates, affecting the quality of the gas produced and potentially leading to incomplete vaporization of the fluid.

[0064] Suitably, the pressure in the buffer tank is maintained below the saturation vapor pressure of the feed gas at the temperature setpoint of the buffer tank. This aspect of the process is crucial for preventing condensation and ensuring optimal gas delivery characteristics. The dew point is the temperature at which a gas begins to condense into liquid. For a gas, this is the temperature at which the partial pressure of the evaporated component equals the saturation vapor pressure of the respective component, leading to the formation of liquid droplets. Saturation vapor pressure is closely related to the dew point but refers specifically to the pressure at which this condensation occurs. It is influenced by the composition of the gas and the temperature. As temperature increases, the dew point (and thus saturation vapor pressure) for a given gas mixture increases, meaning that gases can remain in vapor form at higher pressures as long as their temperature is managed. By ensuring that the pressure in the buffer tank remains below the saturation vapor pressure, the system is designed to prevent the formation of liquid droplets. This is important in applications where the presence of liquid phases could impede gas flow, damage equipment, or lead to inefficient operation.

[0065] The system may need to incorporate dynamic adjustments based on real-time temperature and pressure readings. If the temperature fluctuates, system controls may need to compensate to ensure that the buffer tank pressure remains below the saturation vapor pressure.

[0066] In one embodiment, the buffer tank is temperature-controlled.

[0067] In one embodiment, the process involves sensing the temperature of the feed gas upstream of the buffer tank and controlling the energy input in the heat exchanger based on the temperature sensed in the feed gas.

[0068] In one embodiment, the process of the invention is a process for a gas phase allothermic catalytic conversion, comprising

[0069] I) providing a reactor having a first chamber, a reaction zone in fluid communication with the first chamber and a second chamber in fluid communication with the reaction zone, the reaction zone containing a catalyst material and being electrically heated, each of the first and second chambers containing a solid heat exchange material and being selectively operable in a cooling and a heating mode;

[0070] II) introducing the feed gas from the buffer tank into a selected one of the first and second chambers when the selected chamber is in heating mode and the other of the first and second chamber is in cooling mode; conducting the feed gas through the reaction zone so that the feed gas contacts the catalyst material and forms a gaseous reaction product; and conducting the gaseous reaction product through the other of the first and second chamber;

[0071] III) stopping the flow of feed gas into the selected one of the first and second chambers, and introducing a purge gas, preferably gaseous reaction product, into the selected chamber, thereby purging the selected chamber and pushing feed gas which has remained in the selected chamber through the reaction zone;

[0072] IV) repeating step II) while reversing the direction of feed gas flow;

[0073] V) repeating step III) while reversing the direction of purge gas flow;

[0074] VI) repeating steps II) to V) at least once so that the first chamber and the second chamber alternately operate in the cooling and heating modes during step II).

[0075] The first and second chambers of the reactor contain a solid heat exchange material. The heat exchange material acts as regenerative heat exchangers and has heat retention and heat exchanging properties. Heat exchange materials should be capable of withstanding process temperatures and pressures, chemically inert with respect to the components of both the feed gas and gaseous reaction product mixture. Examples of suitable heat exchange materials which may be used include ceramic or refractory materials such as alumina, silica, zirconia and mixtures thereof.

[0076] For optimization of the heat transfer, the heat exchange material may exhibit one or more or all of the following properties:

[0077] Suitably, the heat exchange material has a density of 500 to 20,000 kg / m3, preferably 1 ,000 to 15,000 kg / m3, more preferably 1 ,500 to 8,000 kg / m3.

[0078] Suitably, the heat exchange material has a heat capacity of 100 to 2,500 J / kg / K, preferably 200 to 2,000 J / kg / K, more preferably 300 to 1 ,500 J / kg / K.

[0079] Suitably, the heat exchange material has a heat conductivity 1 to 30 W / m / K, preferably 1 to 20 W / m / K, more preferably 1 to 10 W / m / K.

[0080] The solid heat exchange materials forms a gas-permeable bed or packing. The solid heat exchange materials may comprise packing material such as beads or spheres, ceramic honeycomb materials, ceramic tubes, and the like. Preferably, the heat exchange material comprises a structured packing, preferably honeycomb monoliths, or comprised of stacked layers of tiles or corrugated materials. Monolithic materials tend to be provided as larger blocks of material, allowing easy and consistent orientation of intrinsic geometric features when packed. The honeycomb monoliths may have straight channels to minimize pressure drop and enable greater reactor length. Alternatively, packing may be more tortuous.

[0081] For optimization of the heat transfer, the bed or packing of the heat exchange material may exhibit one or more or all of the following properties:

[0082] Suitably, the bed or packing of the heat exchange material has a heat capacity of 600 to 5,000 J / m3 / K, preferably 900 to 4,000 J / m3 / K.

[0083] Suitably, the bed or packing of the heat exchange material has a packing density (defined as the ratio of the mass of solid by the volume occupied by the solid and voids) of 250 to 10,000 kg / m3, preferably 500 to 7,500 kg / m3, more preferably 750 to 5,000 kg / m3.

[0084] Suitably, the bed or packing of the heat exchange material has a void fraction (defined as the ratio of void space to the volume occupied by the by the solid and voids) of 0.1 to 0.95, preferably 0.2 to 0.85, more preferably 0.3 to 0.75, most preferably 0.3 to 0.5. Suitably, the bed or packing of the heat exchange material has a specific surface area (defined as the total surface area of material per unit volume) 150 to 5,000 m2 / m3, preferably 300 to 4,000 m2 / m3, more preferably 600 to 3,000 m2 / m3.

[0085] Suitably, flow channels defined by the bed or packing of the heat exchange material can have a hydraulic diameter of 0.5 to 20 mm, preferably 0.5 to 10 mm, more preferably 0.5 to 5 mm.

[0086] The reactor comprises a reaction zone which is in fluid communication with each of the first and second chambers. The reaction zone and the two chambers may be disposed in substantially linear alignment, e.g., vertical or horizontal alignment. Alternatively, the reaction zone and the two chambers may be disposed in pi-shape or inverted pi-shape with a crossbar forming the top of the pi-shape housing the reaction zone, and the first leg and the second leg forming the descenders of the pi-shape and housing the first and second chamber. In still further embodiments, the reaction zone and the two chambers may be disposed in U shape configuration, inverted U shape configuration with the reaction zone forming an arc and the two chambers extending upwardly or downwardly from opposite sides of the reaction zone. The reaction zone contains a catalyst material and is electrically heated.

[0087] The shape of the reaction zone is not especially limited and the reaction zone may, for example, be tubular. Since however the innermost region of a cylindrical solid catalyst arrangement is less efficiently electrically heated than the outer region, a reaction zone having an essentially annular cross section is generally preferred. A reaction zone having annular cross section may be formed by an outer tube and an essentially coaxial inner displacer.

[0088] The use of electric energy as a heat source instead of heating by combustion of natural gas allows considerable advantages, in particular with regard to the ease of control. The electric energy is transferred almost exclusively to the al lothermic reaction with reduced loss of waste heat by combustion off gases. The use of electricity offers opportunities for the use of compact, modular, high performance and energy efficient reactors. When the electricity comes from a non-fossil resource, the allothermic reaction can be implemented without net emission of carbon dioxide.

[0089] There are several ways of directly using electricity as an energy source for carrying out thermochemical reactions.

[0090] The electric heating may be selected from microwave heating, plasma heating, resistive heating selected from direct electric heating of the catalyst material, direct electric heating of a support material, indirect resistive heating, induction heating or combinations of two or more thereof. In particular, the electric heating is by resistive heating, in particular resistive heating selected from

[0091] - heating by resistively heatable heating elements which are in heat-transferring relationship with the catalyst material; and - direct electric heating of the catalyst material, direct electric heating of a support material, or combinations thereof.

[0092] Generally, inductive heating is the process of heating an electrically conductive object (also referred to as a "susceptor") by magnetic induction, through heat generated in the object by eddy currents (also called Foucault currents) and / or by hysteresis loss.

[0093] For example, a catalyst material may comprise electrically conductive catalyst particles or catalyst particles (that may or may not be electrically conductive) mixed with an electrically conductive inert support material, so as to act as a susceptor. Alternatively, inductively heatable heating elements incorporating a susceptor may be immersed in the catalyst material or may otherwise be in direct physical and thermal contact with the catalyst material.

[0094] A preferred susceptor incorporated in an inductively heatable heating element may comprise or consist of a ferromagnetic material, for example a ferromagnetic alloy. Metallic susceptors may specifically be made of at least one material selected from the group consisting of: an iron-based alloy and a nickel-based alloy. Metallic susceptors may specifically be made of at least one ferritic iron-chromium-aluminum alloy (FeCrAI alloy) or one alloy having a material number according to DIN 17007-2: 1961-09: n1.m1 m2m3m4. n1 may specifically be a digit from the group 1 ,2,3 and preferably a digit from the group 1 and 2. ml may specifically be a digit from the group 0, 1 , 3, 4, 8, preferably a digit from the group 3 and 4 and particularly preferably the digit 4. m2 may specifically be a digit from the group 0, 1 , 2, 3, 4, 5, 6, 7, 8, 9, preferably a digit from the group 5, 6, 7, 8, 9 and particularly preferably a digit from the group 7, 8, 9. m3 and m4 preferably each independently may stand for one of the digits from the group 0, 1 , 2, 3, 4, 5, 6, 7, 8, 9. FeCrAI alloys comprise for example products distributed under the trade name Kanthal® AF, Kanthal® A-1 , Kanthal® D.

[0095] For induction heating, the reaction zone may be surrounded by an induction coil energized by a power source supplying an alternating current.

[0096] Alternatively, if the catalyst material is electrically conductive or comprises electrically conductive material, heating can be accomplished by direct resistive heating, namely by passing an electrical current through the catalyst material via electrodes that are in an electrically conductive relationship with the catalyst material. For example, a catalyst material may comprise electrically conductive catalyst particles or catalyst particles (that may or may not be electrically conductive) mixed with an electrically conductive inert support material, electrical current may be passed through the catalyst bed by electrodes immersed in the catalyst material. The resistance of the catalyst material to the electrical current through the electrodes results in heat being liberated within the catalyst material.

[0097] Alternatively, resistively heatable heating elements are used which are in heat-transferring relationship with the catalyst material. Such heating element converts electricity into heat through the process of Joule heating. Electric current through the element encounters resistance, resulting in heating of the element. Heating elements can take the shape of a wire, ribbon, sheet or strip and can be straight, meandering or coiled. Suitable heating elements include heating cables with an electrically insulating mineral shell, tubular heating elements, cartridge heating elements, heating pannels, heating bars, heating plates and heating rings.

[0098] In an embodiment, sheet-shaped heating elements are used that are bent to a cylindrical arc-shape. The cylindrical arc-shaped heating elements are coaxially arranged to form annular gaps in which catalyst is located. The sheet-shaped heating elements preferably comprise a heater strip arranged in a meandering pattern. The cylindrical arc-shaped heating elements are preferably connected to electrodes in a star-shaped pattern, with counter electrodes spaced apart in the flow direction of the fluid process stream, respectively.

[0099] Metallic resistors may specifically be made of at least one material selected from the group consisting of: an iron-based alloy, a nickel-based alloy, a platinum group metal (PGM) or a refractory metal. Metallic resistors may specifically be made of at least one ferritic iron-chromium-aluminum alloy (FeCrAI alloy) or one alloy having a material number according to DIN 17007-2:1961-09: n1.m1m2m3m4. n1 may specifically be a digit from the group 1 ,2,3 and preferably a digit from the group 1 and 2. ml may specifically be a digit from the group 0, 1 , 3, 4, 8, preferably a digit from the group 3 and 4 and particularly preferably the digit 4. m2 may specifically be a digit from the group 0, 1 , 2, 3, 4, 5, 6, 7, 8, 9, preferably a digit from the group 5, 6, 7, 8, 9 and particularly preferably a digit from the group 7, 8, 9. m3 and m4 preferably each independently may stand for one of the digits from the group 0, 1 , 2, 3, 4, 5, 6, 7, 8, 9. FeCrAI alloys comprise for example products distributed under the trade name Kanthal® AF, Kanthal® A-1 , Kanthal® D. PGM comprise the elements platinum, palladium, iridium, rhodium, osmium, ruthenium and alloys thereof. Refractory metals comprise the elements titanium, zirconium, hafnium, niobium, molybdenum, tantalum, tungsten, rhenium and alloys thereof.

[0100] Tungsten, molybdenum, a silver-palladium alloy, silver, a nickel-chromium alloy, an iron-nickel-alloy, an iron- nickel-chromium-alloy or an iron-nickel-chromium-aluminum alloy are preferably used as metallic resistors for the heating elements.

[0101] As an alternative to metallic materials, electrically conductive silicon-based materials and / or carbon-based materials can also be used, particularly preferably molybdenum disilicide or silicon carbide.

[0102] Alternatively, plasma heating may be used. Plasma may be obtained by converting at least part of a gaseous stream into a plasma by partial ionization. The energy released during the recombination of electrons and ions is largely transferred to the gas atoms or molecules, so that a heated process stream is generated which is either itself already a heated reactant stream and / or acts as a heat carrier whose heat is subsequently transferred in direct contact and / or via a heat exchanger to a reactant stream of the heatconsuming process. Conveniently, at least a portion of the gaseous process streams is passed through a plasma generator, into which a cold gas stream is introduced and from which a heated gas stream is withdrawn. The size of the partial quantity that is passed through the plasma generator is determined by the amount of heat to be transferred.

[0103] Plasma heating may be performed using a plasma source selected from a DC plasma, AC plasma, RF plasma or microwave plasma. Preferably, the plasma source has a high specific enthalpy, thus reducing the number of individual heat sources needed to obtain a desired production capacity. The specific enthalpy of the plasma is preferably from 5 MJ / kg to 200 MJ / kg, more preferably from 5 MJ / K to 100 MJ / K.

[0104] Generally, the temperature in the reaction zone is from 200 to 2,500 °C, preferably 400 to 1 ,500 °C, more preferably 600 to 1 ,000 °C.

[0105] Suitably, the reactor is provided with a heat insulating lining. The heat insulating lining is preferably made of a ceramic or refractory material to withstand the high operating temperatures. In general, insulation grade ceramics tend to be relatively porous. Porous ceramics have many open or closed internal pores that provide the thermal barrier properties.

[0106] The reactor includes an outer pressure jacket. For ease of assembly and disassembly, the pressure jacket is preferably split in sections with flanged ends that may be secured by bolts or the like to the flange of an adjacent section of the pressure jacket or the flange of the terminal hoods of the reactor. Preferably, the reaction zone corresponds to a section of the pressure jacket so that the catalyst is accessible from the front side when the jacket is disassembled. The pressure jacket may have pressure-tight passages for the electrodes that provide electricity to the heating elements within the reaction zone.

[0107] The time period after which the direction of gas flow through the reactor is reversed depends on the particular allothermic catalytic conversion that is performed. Generally, the direction of gas flow through the reactor is reversed after a time period of 10 s to 2 h, preferably 30 s to 30 min, more preferably 2 min to 20 min.

[0108] The reactor and method of the invention may be applied to a wide variety of allothermic catalytic conversion. More particularly, the allothermic catalytic conversion may be selected from

[0109] (i) reforming of ammonia according to the idealized equation

[0110] 2 NH3-> N2+ 3H2e.g., catalyzed by transition metal (such as Fe, Ni, Co, Rh or Ru) supported on a refractory support material;

[0111] (ii) Steam cracking of light hydrocarbons,

[0112] (iii) Reverse Water Gas Shift Reaction according to the idealized equation

[0113] CO2+ H2-> CO + H2O e.g., catalyzed by Ni catalyst; (iv) Methane Steam Reforming according to the idealized equation

[0114] CH4+ H2O ^ CO + 3 H2e.g., catalyzed by SIC, C, Ni catalyst or Fe catalyst;

[0115] (v) Methane Dry Reforming according to the idealized equation CH4+ CO2-> 2 CO + 2 H2e.g., catalyzed by SIC, C, Ni catalyst or Fe catalyst;

[0116] (vi) styrene synthesis according to the idealized equation CsHio CsHs + H2e.g., catalyzed by FeO / Pt catalyst;

[0117] (vii) alkane dehydrogenation according to the idealized equation CnH(2n+2) CnH(2n) + H2, wherein n = 2, 3, 4 e.g., catalyzed by FeO / Pt catalyst;

[0118] (viii) cyclohexane dehydrogenation according to the idealized equation C6Hl2 ^ C6H6 + 3 H2e.g., catalyzed by Pt catalyst;

[0119] (lx) dehydrogenation of n-butene according to the idealized equation (C4H8C4H6+ H2),

[0120] (x) alcohol dehydration according to the idealized equation CnH(2n+i)OH -> CnH(2n) + H2O, wherein n = 2, 3, 4 e.g., catalyzed by zeolith catalyst;

[0121] (xi) reforming of alcohols according to the idealized equation e.g., catalyzed by Cu catalyst;

[0122] (xii) decomposition of methanol according to the idealized equation CH3OH -> CO + 2 H2e.g., catalyzed by Cu catalyst;

[0123] (xiii) formation of hydrocyanic acid by the reaction of ammonia with hydrocarbons according to the idealized equation

[0124] CH4+ NH3^ HCN + 3 H2e.g., catalyzed by carbon catalyst;

[0125] (xiv) formation of hydrocyanic acid by formamide cleavage according to the idealized equation HCONH2^ HCN + H2O e.g., catalyzed by stain less particles or Fe catalyst;

[0126] (xv) Coupling of methane according to the idealized equation CH4C2H4, C3H6, C3H6, C3H8 (xvi) dehydroaromatization of methane according to the idealized equation

[0127] 6 CH4C6H6+ 9 H2e.g., catalyzed by Mo catalyst;

[0128] (xvii) formation of anhydrous formaldehyde according to the idealized equation CH3OH -> CH2O + H2e.g., catalyzed by Ag catalyst;

[0129] (xviii) vinyl formamide synthesis from cyanoethyl formamide according to the idealized equation CH3CH(CN)(NH-COH) CH2CH(NH-COH) + HCN e.g., catalyzed by potassium hydroxide on alumina;

[0130] (xix) melamine synthesis according to the idealized equation

[0131] 6 (NH2)2CO C3N6H6 + 6 NH3+ 3 CO2e.g., catalyzed by bauxite.

[0132] In a preferred embodiment, the allothermic catalytic conversion is the reforming of ammonia, where the feed gas comprises ammonia and the gaseous reaction product comprises hydrogen and nitrogen. The hydrogen may then be separated from the gaseous reaction product, for example by a membrane process or a pressure swing adsorption process.

[0133] Hence, the present invention includes a process for recovering hydrogen from ammonia, which comprises pressurizing liquid ammonia, vaporizing ammonia, reforming the ammonia feed gas to a hydrogencontaining gaseous reaction product, and separating hydrogen from the gaseous reaction product.

[0134] In the process of the invention, liquid ammonia is taken from a tank, advantageously at -35 to 50 °C and 1 to 20 bar and, if necessary, brought to higher pressures by means of a pump. The liquid ammonia advantageously becomes the gaseous reformer feed in the evaporator at pressures between 2 and 110 bar, which are the same throughout the process after adjustment for pressure loss. Advantageously, the pressure in the evaporator is between 4 and 113 bar, particularly preferably between 6 and 80 bar, especially between 10 and 70 bar. From the pressure data, the skilled person can determine the temperatures required for evaporation, advantageously -20 to 132 °C, particularly 0 to 110 °C, especially 25 to 100 °C.

[0135] The NH3vapor stream is fed to the reactor where it is split into H2and N2. The energy required for splitting is provided by the electrical heating. The ammonia reforming advantageously takes place at temperatures of 400 to 1 ,100 °C, preferably 600 to 1 ,000 °C, more preferred 700 to 900 °C. Advantageously, the ammonia reforming takes place at a pressure of 4 and 100 bar, particularly preferably between 6 and 80 bar, especially between 10 and 70 bar .

[0136] The catalytic material may be any solid material that catalyzes the reaction. The catalytic material may comprise a supported catalyst or a non-supported catalyst. Preferably, the catalytic material comprises a transition metal supported on a refractory support material. Preferably, the transition metal is selected from the group consisting of Fe, Cu, Ni, Co, Ru, Ag, Pd, Rh, Pt, Ir, including combinations of two or more thereof, more preferably selected from the group consisting of Fe, Ni, Co, Rh and Ru, including combinations of two or more thereof. More preferably, the transition metal is Ni, Co or Ru.

[0137] Catalysts for reforming ammonia are described in the prior art (see A. Di Carlo, et.al, "Ammonia decomposition over commercial RU / AI2O3 catalyst: An experimental evaluation at different operational pressures and temperatures”, International Journal of Hydrogen Energy, 39 (2014) pp. 808-814). For example, ruthenium is used as active catalyst components, advantageously ACTA Hypermec 10010 catalyst (RU / AI2O3).

[0138] The conversion of the ammonia reforming is generally 50 to 99.99%, preferably 80 to 99.9%, particularly preferably 85 to 99%.

[0139] The gaseous ammonia reformate comprises H2, N2 and unreacted NH3. The preferred composition of the gaseous ammonia reformate is as follows: 50 to 75 vol.-% H2, 15 to 25 vol.-% N2, 0 to 30 vol.-% NH3.

[0140] The hydrogen may then be separated from the gaseous ammonia reformate by processes customary in the art, such as a membrane process or pressure swing adsorption process.

[0141] In the membrane process, supported membranes are typically used for hydrogen separation and purification. The supported membrane is made by deposition of a thin layer of metallic membrane on a porous support.

[0142] As the porous support, ceramics or stainless steel supports may be used. Palladium or palladium alloys such as Pd-Ag or Pd-Ag-Au, for example, Pd with 20 to 30 wt.-% Ag, especially with 23 to 24 wt.-% Ag, are used as the metallic membrane. See A. Unemoto, A. Kaimai, S. Kazuhisa, T. Otake, K. Yashiro, J. Mizusaki, T. Kawada, T. Tsuneki, Y. Shirasaki, and I. Yasuda, "The effect of co-existing gases from process of steam reforming reaction on hydrogen permeability of palladium alloy membrane at high temperatures”, International Journal of Hydrogen Energy, No. 32, pp. 2881-2887, 2007.

[0143] The metallic membrane layer thicknesses are preferably between 1 and 60 pm, particularly preferably between 2 and 40 pm, especially between 4 and 20 pm.

[0144] In principle, all known designs can be considered as membrane modules. Among flat membranes, plate modules are a preferred design. As tubular membranes, capillary modules are preferred in addition to hollow fiber modules. Tubular modules with diameters of 1 to 20 mm diameter are particularly preferred, especially with diameters of 2 to 10 mm. The temperature range for H2 separation with membranes, advantageously Palladium- or Palladium-alloy membranes, is advantageously between 250 and 600 °C, particularly preferably between 300 and 500 °C and especially between 350 and 450 °C. Optionally, the temperature of the ammonia reformate is controlled by a heat exchanger between 250 and 600 °C, particularly preferably between 300 and 500 °C and especially between 350 and 450 °C for H2 separation.

[0145] The H2 flux across the membrane is advantageously between 0.1 and 5.0 mol H2 / (m2s), preferably between 0.5 and 4.0 mol H2 / (m2s), particularly preferably between 1.0 and 3.5 mol H2 / (m2s), especially between 1.5 and 3.0 mol H2 / (m2s).

[0146] In the membrane module, the gaseous reformate is split into a high-purity hot permeate stream, with a purity of preferably more than 99.99 vol.-% H2, and a retentate stream, which contains unreacted NH3 in addition to the N2 and H2.

[0147] The retentate preferably includes the following gas composition: 5 to 35 vol.-%-H2, 1 to 40 vol.-% NH3, 25 to 94 vol.-% N2, particularly preferably 10 to 25 vol.-% H2, 5 to 30 vol.-% NH3, and 45 to 85 vol.-% N2.

[0148] The H2 content of the permeate is advantageously 95 to 99.9999 vol.-% H2, particularly preferably 98 to 99.9995 vol.-% H2, especially 99.0 to 99.999 vol.-% H2. The absolute pressure of the permeate is advantageously between 0.1 and 5 bar, particularly preferably between 0.5 and 3.0 bar, especially between 1.0 and 2.0 bar.

[0149] On the permeate side, water vapor may be used as a dilution gas for H2. The water vapor lowers the H2 partial pressure on the permeate side. This increases the driving pressure difference and the H2 flux.

[0150] It may be useful to ensure or further increase the purity of the permeate by passing it over an adsorber bed which separates the last remnants of N2 and NH3 from the permeate. This adsorbent bed then operates as a "police filter”.

[0151] The sensible heat of the permeate of the membrane process can be exploited. For example, the permeate of the membrane process may be heat-exchanged with liquid ammonia to preheat the ammonia prior to vaporization to provide the feed gas.

[0152] Preferably, the retentate is exploited for its calorific value. To this end, the retentate may be combusted with ambient air and the combustion gas may be passed through one or more heat exchangers. The burner burns the combustible components in the retentate, in particular (residual) ammonia and hydrogen, by means of air. The mixture of retentate and heated air is then burned in a burner, e.g. an atmospheric burner or catalytic burner. The combustion gas may be heat-exchanged with the gaseous reaction product leaving the reactor before the gaseous reaction product is directed to membrane separation, and the combustion gas may be subsequently heat-exchanged with ammonia to vaporize the ammonia and provide the feed gas.

[0153] The air required for the burner is advantageously drawn in from the environment. The air drawn in is then advantageously compressed for conveying the hot combustion gas via the heat exchangers. Advantageously, the air is compressed from ambient pressure (1.013 bar) to 1.05 to 5.0 bar, preferably to 1.1 to 2.0 bar, in particular 1.2 to 1.5 bar. All devices known to the person skilled in the art can be used as compressors, e.g. fans, ventilators, compressors etc. The compressor is advantageously located upstream of the first burner.

[0154] Using an atmospheric burner, the hot combustion gas produced in the burner advantageously has a temperature of 500 to 1 ,200 °C or 600 to 1 ,100 °C, preferably 700 to 1 ,000 °C, particularly preferably 800 to 950 °C, especially 850 to 900 °C. Using a catalytic burner, it advantageously has a temperature of 200 to 700 °C, preferably 220 to 500 °C.

[0155] The combustion gas advantageously includes N2, O2 and H2O. Exemplarily, the fuel gas has the following composition: 80 vol.-% N2, 10 vol.-% O2, and 10 vol.-% H2O.

[0156] The combustion gas is preferably heat-exchanged with the gaseous reaction product leaving the reactor before it is directed to membrane separation, and subsequently heat-exchanged with ammonia to vaporize the ammonia and provide the feed gas. To this end, the combustion gas is passed through successive heat exchangers.

[0157] Pressure swing adsorption involves a pressure swing adsorption (PSA) unit, which consists of a series of beds filled with adsorbent material, for example zeolites or activated carbons, in particular carbon molecular sieves. As the pressurized reformate flows through the bed gaseous species adsorb on to the active surfaces. Since the H2 is the least strongly adsorbed species in the reformate stream, a pure H2 gas exits the bed. After a period of time when the adsorbent sites begin to become saturated, the feed gas is removed and the bed is depressurized forcing the adsorbed species to desorb and exit the bed as the off-gas or tail gas stream. By cycling several beds through this pressurization and depressurization cycle a continuous H2 purification process is created. The pressure-swing adsorption (PSA) device can also function as a vacuumswing adsorption (VSA) device with the aid of a vacuum pump. Useful PSA devices are described in US 20230242395 A1 and KR 102247199 B1.

[0158] Utilizing a PSA device, the gaseous reaction product from the reactor is subjected to a cooling process using a heat exchanger, preferably at a temperature below 100 °C, particularly within the range of 0 to 60 °C. The cooled gas is then directed to a pressure-swing adsorption (PSA) device or vacuum swing adsorption (VSA) device. Within the PSA device, the separation of nitrogen (N2) and ammonia (NH3) from hydrogen (H2) takes place. The PSA device can conveniently operate at the outlet pressure of the reactor, resulting in the production of pure H2 at a high pressure. The tail gas from the PSA, consisting of N2, N H3, and H2, is released at a low pressure and may be exploited for its calorific value. To this end, the tail gas may be combusted with ambient air and the combustion gas may be passed through one or more heat exchangers in essentially the same way as described above with respect to the retentate. In a preferred embodiment, the ambient air is preheated by heat exchange with the gaseous reaction product.

[0159] In another embodiment the tail gas from the PSA is directed to a TSA unit to remove NH3. During the regeneration of the TSA a hot N2 stream is used to desorb NH3. The hot N2 stream is generated by heating N2 with a heat exchanger in contact with the hot off-gas stream from the incinerator.

[0160] In another embodiment, NH3 is removed from the gaseous reaction product before entering the PSA device using a temperature-swing adsorption (TSA) device operating at high pressure. The PSA device subsequently separates N2 from H2, while the byproduct gas is directed to the incinerator. During the regeneration of the TSA a hot N2 stream is used to desorb NH3. The hot N2 stream is generated by heating N2 with a heat exchanger in contact with the gaseous reaction product from the reactor or with the hot offgas stream from the incinerator.

[0161] The invention is further illustrated by the appended drawings.

[0162] Fig. 1 schematically shows an electrically heated reverse flow reactor with a linear arrangement of the reaction zone und first and second chamber and side draws in the first and second chambers.

[0163] Fig 2. schematically shows an electrically heated reverse flow reactor with an U-shaped arrangement of the reaction zone und first and second chamber.

[0164] Fig. 3A schematically shows a longitudinal cross-section of a reactor according to the invention.

[0165] Fig. 3B schematically shows a traverse cross-section of the reaction zone of a reactor according to the invention.

[0166] Fig. 4 is a flow diagram of an inventive process for reforming ammonia including an electrically heated reverse flow reactor, hydrogen separation via a membrane process, and various units for heat integration.

[0167] Fig. 5 shows a flow diagram of an inventive process for reforming ammonia including an electrically heated reverse flow reactor with a side withdrawal, hydrogen separation via pressure swing adsorption, and various units for heat integration including a step for the oxidation of the tail gas of the PSA unit. Fig. 6 shows a flow diagram of an inventive process for reforming ammonia including an electrically heated reverse flow reactor with a side withdrawal, hydrogen separation via pressure swing adsorption, and various units for heat integration including a step for the oxidation of the tail gas of the PSA unit. The calorific value of the reactor product extracted from the side draw is at least partially utilized for the generation of high pressure steam at a high energy level.

[0168] Fig. 7 shows the mass fraction profile and the temperature profile of a simulation for an exemplary electrically heated reverse flow reactor used for ammonia reforming into hydrogen and nitrogen.

[0169] Fig. 8 shows the mass fraction profile and the temperature profile of a simulation for an exemplary electrically heated reverse flow reactor with a partial side draw used for ammonia reforming into hydrogen and nitrogen.

[0170] Fig. 9 schematically depicts the cycling operational steps underlying an embodiment of the process of the invention for the above-described embodiment of a gas phase catalytic conversion, wherein the process comprises active and inactive phases.

[0171] Fig. 10 schematically depicts a plant comprising an evaporator, a buffer tank and a reactor.

[0172] Fig. 11 A depicts an ammonia feed flow rate to the reactor over time in an exemplary process of reforming of ammonia, wherein the process comprises active and inactive phases. Fig. 11 B depicts pressure changes in the buffer tank over time in an exemplary manner for the process of the invention, wherein the process comprises active and inactive phases.

[0173] According to Fig. 1 , reactor 101 has two chambers 102, 104 containing a solid heat exchange material, and a reaction zone 103 sandwiched between the chambers 102, 104 in a linear configuration. The reaction 103 zone comprises a catalyst material and is heated by heater 110.

[0174] Valves 105A and 108B are open when the first chamber 102 is in the heating mode and the second chamber 104 is in the cooling mode, while valves 105B and 108A are closed. Feed gas is introduced into the first chamber 102 via line 106 and valve 105A, conducted through the reaction zone 103 and the the second chamber 104, and gaseous reaction product leaves the reactor via valve 108B and line 109. To reverse the direction of gas flow through the reactor valves 105B and 108A are opend while 105A and 108B are closed. A portion of the hot gaseous reaction product is discharged via one of valves 106A, 106B from the chamber operating in cooling mode, and utilized further via line 107.

[0175] According to Fig. 2, inverted U-shaped reactor 201 has two chambers 202, 204 containing a solid heat exchange material, and a reaction zone 203 arranged in an "arch" between the chambers 202, 204. The reaction 203 zone comprises a catalyst material and is heated by heater 210. Valves 205A and 208B are open when the first chamber 202 is in the heating mode and the second chamber 204 is in the cooling mode, while valves 205B and 208A are closed. Feed gas is introduced into the first chamber 202 via line 206 and valve 205A, conducted through the reaction zone 203 and the the second chamber 204, and gaseous reaction product leaves the reactor via valve 208B and line 209. To reverse the direction of gas flow through the reactor valves 205B and 208A are opend while 205A and 208B are closed. A portion of the hot gaseous reaction product is discharged via one of valves 206A, 206B from the chamber operating in cooling mode, and utilized further via line 207.

[0176] According to Fig. 3A, reactor 301 has two chambers 302, 304 containing a solid heat exchange material 305, 306, and a reaction zone 303 sandwiched between the chambers 302, 304. Electrodes 307, 308 protrude into the reaction zone 303, which supply electricity to heating elements 309 extending between electrodes 307, 308. The gaps between heating elements 309 contain a catalyst material (not shown in Fig. 3A). The heat-insulating lining 311 together with displacer 310 forms an annular reaction zone.

[0177] According to Fig. 3B, three electrodes 307A, 307B and 307C are arranged in a star-shaped pattern and are connected to arc-shaped heating elements 309. The gaps between heating elements 309 contain a catalyst material (not shown in Fig. 3B). The heat-insulating lining 311 together with displacer 310 forms an annular reaction zone.

[0178] According to Fig. 4, liquid ammonia is supplied via line 405 to an ammonia storage tank 406. Liquid ammonia is withdrawn from tank 406 via pump 407 and directed to ammonia preheater 408. Preheated ammonia is directed to ammonia vaporizer 409 where the preheated ammonia is vaporized. In the shown position of the four-way valve 402, gaseous ammonia is directed to reactor 401 via line 403. The reactor 401 has two chambers containing a solid heat exchange material, and a reaction zone sandwiched between the chambers containing a catalyst material and being electrically heated (not shown in Fig. 4).

[0179] The gaseous reaction product (reformate) containing H2, N2 and unreacted NH3 leaves the reactor via line 404 and four-way valve 402, and is directed via line 410 and heat exchanger 411 to membrane module 412. A permeate which is essentially pure H2 leaves the membrane module 412 via line 413. The sensible heat of permeate 413 is used in heat exchanger 408 to preheat liquid ammonia. A retentate containing unreacted NH3 in addition to the N2 and H2 leaves the membrane module 412 via line 414 and is directed to incinerator 415, where the retentate is combusted with air that is supplied via line 416. The hot combustion gas obtained from incinerator 415 is used in heat exchanger 411 to adjust the temperature of the reformate from reactor 401 before it enters the membrane module 412, and in heat exchanger 409 to vaporize preheated liquid ammonia.

[0180] According to Fig. 5, reactor 501 has two chambers 511 A, 511 B containing a solid heat exchange material, and a reaction zone 512 sandwiched between the chambers and containing a catalyst material and being electrically heated via a heater 513. Valves 510A and 504B are open when the first chamber 511 A is in the heating mode and the second chamber 511 B is in the cooling mode, while valves 51 OB and 504A are closed.

[0181] Liquid ammonia is supplied via pump 507 and directed to ammonia preheater 508. Preheated ammonia is directed to ammonia vaporizer 509 where the preheated ammonia is vaporized. Gaseous ammonia is introduced into the first chamber 511 A via valve 510A, conducted through the reaction zone 512 and the the second chamber 511 B, and gaseous reaction product leaves the reactor via valve 504B. To reverse the direction of gas flow through the reactor valves 510B and 504A are opend while 510A and 504B are closed.

[0182] A portion of the gaseous reaction product is discharged via one of valves 514A, 514B from the chamber operating in cooling mode. The sensible heat of the discharged portion is used to vaporize preheated ammonia via the heat exchange circuit consisting of ammonia vaporizer 509 and heat exchanger 503, before the discharged portion is combined with the gaseous reaction product emanating from the reactor.

[0183] The gaseous reaction product (reformate) containing H2, N2 and unreacted NH3 leaves the reactor via one of valves 504A, 504B, and is directed via heat exchanger 505 to a pressure swing adsorption (PSA) device 506. Essentially pure H2 leaves the PSA device via line 515. A tail gas containing unreacted NH3 in addition to the N2 and H2 leaves the PSA device via line 516 and is directed to incinerator 517, where the tail gas is combusted with air that is supplied via line 518 and preheated. The hot combustion gas obtained from incinerator 517 is used in heat exchanger 519 to preheat ammonia before it is vaporized.

[0184] According to Fig. 6, reference signs 601 to 604 and 610 to 614 are used to denote the same elements as those having reference signs 501 to 504 and 510 to 514 in Fig. 5 and will not be described in detail here.

[0185] The gaseous reaction product (reformate) containing H2, N2 and unreacted NH3 leaves the reactor via one of valves 604A, 604B, and is directed via heat exchanger 607 to a pressure swing adsorption (PSA) device 608. Essentially pure H2 leaves the PSA device via line 609. A tail gas containing unreacted NH3 in addition to the N2 and H2 leaves the PSA device via line 614 and is directed to incinerator 615, where the tail gas is combusted with air that is supplied via heat exchanger 607 and line 616. The hot combustion gas obtained from incinerator 615 is used in heat exchanger 617 to provide heat for ammonia vaporizer 608.

[0186] A portion of the gaseous reaction product is discharged via one of valves 614A, 614B from the chamber operating in cooling mode. The sensible heat of the discharged portion is used to vaporize a partial stream of feed water via heat exchanger 603 to produce pressurized steam which is discharged via line 606, before the discharged portion is combined with the gaseous reaction product emanating from the reactor. Steam condensate is recycled via line 605.

[0187] Fig. 7 shows the mass fraction profile (solid line = NH3, dashed line H2, dotted line = N2) and the temperature profile of a simulation for an exemplary electrically heated reverse flow reactor used for ammonia reforming into hydrogen and nitrogen. Reaction takes place exclusively in the reaction zone extending from 0.5 to 1 m and containing a catalyst material where NH3 is consumed and H2 and N2 are formed.

[0188] The selected chamber in heating mode extending from 0 to 0.5 m contains preheated solid heat exchange material wherein the feed gas is heated to a desired temperature. Subsequently, the feed gas is passed through the electrically heated reaction zone, and finally through the other chamber extending from 1.0 to 1 ,5 m containing an unheated solid heat exchange material. The hot gaseous reaction product being cooled by the unheated heat exchange material in the other chamber. Fig. 7 reveals that the temperature gaseous reaction product at the distal end of the second chamber (1 .5 m) exceeds the temperature the feed gas at the distal end of the first chamber (0 m).

[0189] Fig. 8 shows the mass fraction profile (solid line = NH3, dashed line H2, dotted line = N2) and the temperature profile of a simulation for an exemplary electrically heated reverse flow reactor with a partial side draw (20 vol.-%) used for ammonia reforming into hydrogen and nitrogen. The temperature at the reactor outlet (1.5 m) is close to the temperature at the reactor inlet (0 m).

[0190] In Fig. 9, cycling operational steps C-l) to C-IV) are depicted, which correspond to steps II) to V) of the process of the invention for the above-described embodiment of a gas phase catalytic conversion:

[0191] - Cycling operational step C-l) involves introducing the feed gas from the buffer tank into a first chamber in heating mode. The second chamber is in cooling mode. The feed gas is conducted through the reaction zone so that the feed gas contacts the catalyst material and forms a gaseous reaction product, which is conducted through the second chamber.

[0192] - Cycling operational step C-l I) involves introducing a purge gas (preferably gaseous reaction product) into the first chamber, thereby purging the first chamber and pushing feed gas which has remained in the first chamber through the reaction zone.

[0193] - Cycling operational step C-lll) involves repeating the steps of cycling operational step I) while reversing the direction of feed gas flow. Thus, cycling operational step C-lll) involves introducing the feed gas from the buffer tank into the second chamber in heating mode. The first chamber is in cooling mode. The feed gas is conducted through the reaction zone so that the feed gas contacts the catalyst material and forms a gaseous reaction product, which is conducted through the first chamber.

[0194] - Cycling operational step C-IV) involves repeating the steps of cycling operational step II) while reversing the direction of purge gas flow. Thus, cycling operational step C-IV) involves introducing a purge gas (preferably gaseous reaction product) into the second chamber, thereby purging the second chamber and pushing feed gas which has remained in the second chamber through the reaction zone.

[0195] According to Fig. 10, a liquid feed, e.g., liquid ammonia, from tank 1001 is provided to preheater 1002 via a pump. The flow rate of liquid ammonia is sensed upstream of preheater 1002, and the pump is adjusted according to the sensed flow rate. The preheated liquid is conducted to evaporator 1003 to obtain a feed gas, e.g., gaseous ammonia. The temperature of preheated liquid is sensed upstream of evaporator 1003, and the temperature in preheater 1002 is controlled based on the temperature sensed in the preheated liquid. The temperature of the feed gas is sensed upstream of buffer tank 1004, and the temperature in evaporator 1003 is controlled based on the temperature sensed in the feed gas. The feed gas is conducted into buffer tank 1004. The feed gas pressure upstream the buffer tank is sensed, and a throttle member arranged in the duct of the flow of feed gas to the buffer tank 1004 is adjusted according to the sensed pressure. Buffer tank 1004 is temperature-controlled.

[0196] The plant further comprises a reactor 1006 which intermittently consumes feed gas. Reactor 1006 cycles between an active phase wherein reactor 1006 consumes feed gas and an inactive phase wherein reactor 1006 does not consume feed gas. The flow rate of feed gas upstream of reactor 1006 is sensed and a throttle member arranged in the duct of the flow of feed gas to reactor 1006 is adjusted according to the sensed flow rate. In inactive phases, the flow of feed gas into the reactor is stopped and purge gas is introduced into the reactor 1006 via line 1005, thereby purging the selected chamber and pushing feed gas which has remained in the selected chamber through the reaction zone. The flow rate of purge gas in line 1005 upstream of reactor 1006 is sensed and a throttle member arranged in the duct of the flow of purge gas to reactor 1006 is adjusted according to the sensed flow rate. During inactive phases, feed gas accumulates in the buffer tank 1004, thereby increasing pressure in the buffer tank 1004. In active phases, pressure is relieved in the buffer tank 1004.

[0197] The plant is in particular configured for a process for a gas phase allothermic catalytic conversion. The reactor 1006 has a first chamber 1007, a temperature-controlled reaction zone 1008 in fluid communication with the first chamber 1007 and a second chamber 1009 in fluid communication with the reaction zone 1008. The reaction zone 1008 contains a catalyst material and is electrically heated. Each of the first and second chambers 1007, 1009 contain a solid heat exchange material and are selectively operable in a cooling and a heating mode.

[0198] Valves 1010A and 1011 B are open when the first chamber 1007 is in the heating mode and the second chamber 1009 is in the cooling mode, while valves 1010B and 1011 A are closed. Feed gas is introduced into the first chamber 1007 via valve 1010A, conducted through the reaction zone 1008 and the second chamber 1009, and gaseous reaction product leaves the reactor via valve 1011 B. To reverse the direction of gas flow through the reactor, valves 1010B and 1011 A are opened while 1010A and 1011 B are closed. The hot gaseous reaction product is discharged. The pressure of the discharged hot gaseous reaction product is sensed and a throttle member arranged in the duct of the flow of discharged hot gaseous reaction product is adjusted according to the sensed pressure.

[0199] In Fig. 11 A, the ammonia feed flow rate over time illustrates the repeated cycling between the active phase wherein the reactor consumes feed gas and the inactive phase wherein the reactor does not consume feed gas. Thus, the ammonia feed flow rate is approximately 1 ,700 Nl / h during the active phase and is switched to 0 Nl / h during the inactive phase. The duration of each half-cycle is 180 s, comprising an active phase of 144 s and an inactive phase of 36 s.

[0200] According to Fig. 11 B, the pressure in the buffer tank, after an initial pressure of about 60 bar, fluctuates between about 53 bar and about 57 bar, illustrating the repeated cycling between the active phase wherein the reactor consumes feed gas and the inactive phase wherein the reactor does not consume feed gas.

[0201] Thus, in inactive phases, feed gas accumulates in the buffer tank until a maximum of about 57 bar is reached. In active phases, pressure is relieved in the buffer tank until a minimum of about 53 bar is reached. The dashed line represents the essentially constant pressure of about 60 bar in the evaporator (and in the feed gas supply line upstream of the buffer tank), whereas the dotted line represents the essentially constant pressure of about 50 bar in the reactor (and in the feed gas flow line from the buffer tank to the reactor).

Claims

Claims1 . Process for a gas phase al I othermic catalytic conversion, comprising a) providing a reactor having a first chamber, a reaction zone in fluid communication with the first chamber and a second chamber in fluid communication with the reaction zone, the reaction zone containing a catalyst material and being electrically heated, each of the first and second chambers containing a solid heat exchange material and being selectively operable in a cooling and a heating mode; b) introducing a feed gas into a selected one of the first and second chambers when the selected chamber is in heating mode and the other of the first and second chamber is in cooling mode; c) conducting the feed gas through the reaction zone so that the feed gas contacts the catalyst material and forms a gaseous reaction product; d) conducting the gaseous reaction product through the other of the first and second chamber; e) periodically reversing the direction of gas flow through the reactor so that the first chamber and the second chamber alternately operate in the cooling and heating modes.

2. Process according to claim 1 , comprising discharging a portion of the gaseous reaction product from the one of the first or second chamber operating in cooling mode at an intermediate point between the reaction zone and the distal end of the chamber.

3. Process according to claim 1 or 2, comprising, prior to reversing the direction of gas flow, stopping the flow of feed gas into the selected one of the first and second chambers, and introducing a purge gas into the selected chamber, thereby purging the selected chamber and pushing feed gas which has remained in the selected chamber through the reaction zone.

4. Process according to claim 3, comprising providing a supply of the feed gas and conducting the feed gas through a buffer tank; wherein when the flow of feed gas into the selected one of the first and second chambers is stopped, feed gas accumulates in the buffer tank, thereby increasing pressure in the buffer tank, and when the flow of feed gas into the other of the first and second chamber is resumed, pressure is relieved in the buffer tank.

5. Process according to any one of the preceding claims, wherein the solid heat exchange materials form a gas-permeable bed or packing, having a heat capacity of 600 to 5000 J / m3 / K, preferably 900 to 4000 J / m3 / K, and / or a void fraction of 0.1 to 0.95, preferably 0.2 to 0.85, more preferably 0.3 to 0.75, most preferably 0.3 to 0.5.

6. Process according to any one of the preceding claims, wherein the heat exchange material comprises a structured packing, preferably ceramic honeycomb materials.

7. Process according to any one of the preceding claims, wherein the temperature in the reaction zone is from 200 to 2500 °C, preferably 400 to 1500 °C, more preferably 600 to 1000 °C.

8. Process according to any one of the preceding claims, wherein the electric heating is by resistive heating, in particular resistive heating selected from- heating by resistively heatable heating elements which are in heat-transferring relationship with the catalyst material; and- direct electric heating of the catalyst material, direct electric heating of a support material, or combinations thereof.

9. Process according to any one of the preceding claims, wherein the al lothermic catalytic conversion is selected from(I) reforming of ammonia according to the idealized equation2 NH3-> N2+ 3H2;(ii) Steam cracking of light hydrocarbons;(ill) Reverse Water Gas Shift Reaction according to the idealized equation CO2+ H2-> CO + H2O;(iv) Methane Steam Reforming according to the idealized equation CH4+ H2O ^ CO + 3 H2;(v) Methane Dry Reforming according to the idealized equation CH4+ CO2-> 2 CO + 2 H2;(vi) styrene synthesis according to the idealized equation CsHio CsHs + H2;(vii) alkane dehydrogenation according to the idealized equation CnH(2n+2) CnH(2n) + H2, wherein n = 2, 3, 4;(viii) cyclohexane dehydrogenation according to the idealized equation CeHi2CeHe + 3 H2;(ix) dehydrogenation of n-butene according to the idealized equation (C4H8C4H6+ H2);(x) alcohol dehydration according to the idealized equationCnH(2n+i)OH -> CnH(2tl) + H2O, wherein n = 2, 3, 4;(xi) reforming of alcohols according to the idealized equation (CnH(2n+i)OH + (n-1) H2O nCO + 2n H2), n=1, 2, 3;(xii) decomposition of methanol according to the idealized equation CH3OH -> CO + 2 H2;(xiii) formation of hydrocyanic acid by the reaction of ammonia with hydrocarbons according to the idealized equationCH4+ NH3^ HCN + 3 H2;(xiv) formation of hydrocyanic acid by formamide cleavage according to the idealized equation HCONH2^> HCN + H2O;(xv) Coupling of methane according to the idealized equation CH4C2H4, C3H6, C3H6, C3H8;(xvi) dehydroaromatization of methane according to the idealized equation6 CH4-> C6H6+ 9 H2;(xvii) formation of anhydrous formaldehyde according to the idealized equation CH3OH CH2O + H2;(xviii) vinyl formamide synthesis from cyanoethyl formamide according to the idealized equation CH3CH(CN)(NH-COH) CH2CH(NH-COH) + HCN;(xix) melamine synthesis according to the idealized equation6 (NH2)2CO C3N6H6+ 6 NH3+ 3 CO2.

10. Process according to claim 9, wherein the allothermic catalytic conversion is reforming of ammonia, wherein the feed gas comprises ammonia, and the gaseous reaction product comprises hydrogen and nitrogen.11 . Process according to claim 10, wherein the catalytic material comprises a transition metal supported on a refractory support material, and wherein the transition metal is preferably selected from the group consisting of Fe, Cu, Ni, Co, Ru, Ag, Pd, Rh, Pt, Ir, including combinations of two or more thereof, more preferably from the group consisting of Fe, Ni, Co, Rh and Ru, including combinations of two or more thereof.

12. Process according to claim 10 or 11, additionally comprising separating hydrogen from the gaseous reaction product by a membrane process or a pressure swing adsorption process.

13. Process according to claim 12, wherein the permeate of the membrane process is heat-exchanged with liquid ammonia to preheat the ammonia.

14. Process according to claim 12 or 13, wherein the gaseous retentate of the membrane process is combusted with ambient air and, preferably, the combustion gas is heat-exchanged with the gaseous reaction product leaving the reactor before the gaseous reaction product is directed to membrane separation, and the combustion gas is subsequently heat-exchanged with ammonia to vaporize the ammonia and provide the feed gas.

15. Process according to claim 14, wherein a tail gas of a pressure swing adsorption process is combusted with ambient air and, preferably, the combustion gas is heat-exchanged with ammonia to preheat and / or vaporize the ammonia to provide the feed gas, and / or the ambient air is preheated by heat exchange with the gaseous reaction product.

16. An apparatus for a gas phase allothermic catalytic conversion, including a reactor having a first chamber, a reaction zone in fluid communication with the first chamber and a second chamber in fluid communication with the reaction zone, the reaction zone containing a catalyst material and being electrically heated, each of the chambers containing a solid heat exchange material and being selectively operable in a cooling and heating mode; the reactor being adapted for introducing a feed gas into a selected one of the first and second chambers when the selected chamber is in heating mode and the other of the first and second chamber is in cooling mode; conducting the feed gas through the reaction zone so that the feed gas contacts the catalyst material and forms a gaseous reaction product; and conducting the gaseous reaction product through the other of the first and second chamber; the apparatus further including gas flow directing means for periodically reversing the direction of gas flow through the reactor so that the first chamber and second chamber alternately operate in the cooling and heating modes.

17. The apparatus according to claim 16, comprising outlet means in fluid communication with at least one of- the reaction zone for discharging a portion of the gaseous reaction product from the reaction zone; and- each of the first and second chamber for discharging a portion of the gaseous reaction product from the one of the first or second chamber operating in cooling mode at an intermediate point between the reaction zone and the distal end of the chamber.

18. The apparatus according to claim 16 or 17, the reactor further being adapted for temporarily stopping the feed gas stream and introducing a purge gas stream into the selected one of the chambers before the direction of gas flow through the reactor is reversed.The apparatus according to claim 18, including a buffer tank for temporarily storing a feed gas and a reactor; the buffer tank being adapted to release the feed gas to the reactor.

Images