Method and apparatus for obtaining high purity hydrogen from methanol or ammonia

Abstract

The present invention relates to a method for obtaining hydrogen from methanol or ammonia, which comprises the steps of: (1) evaporating methanol or ammonia; (2) reforming the methanol or ammonia to form a hydrogen-containing gas mixture; (3) cooling the gaseous product to 25-100°C; (4) separating hydrogen from the cooled gaseous product by a sorption process at a pressure of 1-60 bar and a temperature of 25-100°C; (5) compressing and preheating air in parallel with the first four steps; (6) regenerating the adsorbent with the extract using preheated ambient air; and (7) combusting the extract (tail gas) separated from the adsorbent with air. The combustion gas is passed through at least two different heat exchangers in the direction of the combustion gas flow to (i) first provide reaction heat for reforming the methanol or ammonia, and (ii) subsequently provide evaporation heat for evaporating the feed to the reformer. The reformate preheats ambient air for the regeneration process in a heat exchanger before passing it through the sorption step, the separated hydrogen preheats ambient air for the regeneration process after the sorption process, and / or the combustion gases finally preheat ambient air for the regeneration process as step (iii).

Description

[Technical Field]

[0001] The present invention relates to a method for obtaining hydrogen from methanol or ammonia, which comprises the steps of: (1) evaporating methanol or ammonia; (2) reforming the methanol or ammonia to form a hydrogen-containing gas mixture; (3) cooling the gaseous product to 25-100°C; (4) separating hydrogen from the cooled gaseous product by a sorption process at a pressure of 1-60 bar and a temperature of 25-100°C; (5) compressing and preheating air in parallel with the first four steps; (6) regenerating the adsorbent with the extract using preheated ambient air; and (7) combusting the extract (tail gas) separated from the adsorbent with air. The combustion gas is passed through at least two different heat exchangers in the direction of the combustion gas flow to (i) first provide reaction heat for reforming the methanol or ammonia, and (ii) subsequently provide evaporation heat for evaporating the feed to the reformer. The separated hydrogen preheats the ambient air for the regeneration process following the sorption process.

[0002] Hydrogen offers the desired conditions to become a key factor for future energy supply. The transport sector and industry in particular face the major challenge of becoming more climate-friendly. According to calculations by the Federal Network Agency, Germany alone will need approximately 4.3 million tons of hydrogen per year in 2045 to achieve climate neutrality. This corresponds to approximately 144 TWh of thermal power per year (Approval of the Scenario Framework 2023-2037 / 2045, Federal Network Agency, July 2022, (https: / / www.netzausbau.de / Wissen / Ausbaubedarf / Szenariorahmen / de.html)). For comparison, today's large-scale plants have a maximum annual production capacity of less than 100,000 tons of hydrogen. In addition, feedstocks are almost entirely fossil-based.

[0003] In the future, hydrogen will be produced using renewable energy. Because weather conditions in Germany are unfavorable and land availability is limited, these large quantities of hydrogen will likely have to be imported. Preferred hydrogen carriers for long-distance transport are methanol and ammonia. Methanol can also be in the form of crude condensate, which is a mixture of methanol and water of reaction.

[0004] There is a need for a technology that can recover hydrogen from methanol or ammonia efficiently and at low cost.

[0005] In order for hydrogen to be used as a chemical feedstock for industry or in the transportation sector in fuel cell applications, it must be of very high quality, as impurities will affect catalysts and membranes.

[0006] Hydrogen is currently mainly produced centrally in steam methane reforming (SMR) production units. When production and recycling sites are far from each other, the hydrogen must be highly compressed (up to 350 bar) and, in rare cases, liquefied in order to be transported using suitable transport vehicles to where it is needed, such as hydrogen refueling stations. However, transporting hydrogen by vehicle is uneconomical and environmentally unfriendly, as larger hydrogen refueling stations must be served daily by trucks.

[0007] In parallel with vehicle transport, there are several high-purity hydrogen pipelines. However, in order to be able to supply hydrogen to refueling stations on a large scale, it is necessary to build a separate high-density hydrogen pipeline network similar to the natural gas network. However, such pipeline networks have very high infrastructure costs and require complex approval procedures, which makes their realization in the near future unlikely.

[0008] Water electrolysis has very high electricity demands, which must be met as needed using available grid power due to insufficient storage capacity of H2 at refueling stations and industry. In addition to today's electricity consumption, if the automotive sector and heat generation are purely electrical, grid power demand will nearly double in the future, so not only will wind and solar power capacities have to be expanded many times over, but the electricity grid will also have to be significantly expanded.

[0009] In order to preserve the vast land and marine space required for this, while keeping the negative environmental impact (this applies especially to wind turbines) as low as possible and minimizing it, it would make sense to relocate hydrogen production to regions with much more favorable weather conditions and much larger areas of open land than Germany. In countries with very favorable conditions for renewable energies, such as the MENA (Middle East and North Africa) countries, efforts are already underway to create this demand for renewable energy in the future. One example is NEOM HELIOS in Saudi Arabia, the world's largest green hydrogen / ammonia project.

[0010] The production costs of hydrogen are significantly lower in countries with stronger winds and sunnier skies than Germany. To take advantage of this cost advantage, transportation costs must also be low. This requirement is met by two hydrogen carriers: methanol and, above all, ammonia. So-called liquid organic hydrogen carriers (LOHCs) also meet this requirement to some extent.

[0011] Methanol (MeOH) is a basic chemical that can be produced on a large scale and is an excellent energy carrier due to its high energy density of 19.9 MJ / kg. In contrast to hydrogen, methanol can be transported cost-effectively (O. Machhammer, "Regenerativer Strom aus Deutschland oder e-Fuels aus Chile: Worauf sollte die zukunftige Mobilitat bauen?" ["Renewable electricity from Germany or e-fuels from Chile: What should future mobility be based on?"] Chemie Ingenieur Technik, No. 4, 2021). Existing crude oil transportation infrastructure can be used for transportation.

[0012] Methanol is still primarily used today as a base chemical to produce, for example, formaldehyde, acetic acid, methyl chloride, methyl methacrylate, and methylamine. The energy balance plays a subordinate role in these processes; the added value of the resulting products is essential.

[0013] Ammonia (NH3) is a basic chemical that is produced on a large scale, for example for the production of fertilizer. Ammonia is a good energy carrier, with a mass-based energy density of 18.6 MJ / kg, roughly the same as that of methanol (MeOH), at 19.9 MJ / kg. Ammonia has a boiling point of -33°C and can be transported at ambient temperature in low-pressure vessels at 10 bar.

[0014] A key feature of future energy carriers is their low carbon dioxide footprint. In the case of NH3, in addition to renewable produced hydrogen (H2), nitrogen (N2) is also required, which is present in high concentrations in the atmosphere, around 80%, and can therefore be easily extracted via air separation plants.

[0015] In today's main focus, for example the material utilization of ammonia as a fertilizer, the energy balance plays a subordinate role. The fertilizer effect is important in this context.

[0016] Known methods for separating N2 and H2 are distillation processes, sorption processes such as pressure swing adsorption (PSA) or temperature swing adsorption (TSA) or a combination of PSA and TSA, and membrane processes.

[0017] Hydrogen can be made available at refueling stations, for example, for refueling fuel cell (FC) vehicles. For this purpose, hydrogen is compressed to the required pressure of 950 bar for intermediate storage and cooled to the required temperature of -40°C during refueling.

[0018] However, when methanol or ammonia is used as the energy carrier, the energy balance of the overall process plays an important role: the overall process from methanol or ammonia reforming to the release of H should advantageously have low energy losses in order to retain as much of the energy originally used as possible.

[0019] Fuel cell (FC) operation requires hydrogen with very high purity (>99.99%). The production of hydrogen with the highest purity from methanol or ammonia requires several process steps: evaporation and cleavage of the methanol or ammonia, and separation of the high-purity hydrogen from the resulting gas mixture. The thermal energy required for evaporation and cleavage must be supplied externally or provided by burning a portion of the methanol used, the ammonia used, or a portion of the reforming products.

[0020] The prior art primarily focuses on maximum conversion and optimized hydrogen separation in the reforming process.

[0021] Membrane reactors are currently the primary approach to maximize conversion, and for optimized hydrogen separation membrane processes, or a combination of PSA (pressure swing adsorption) and membrane processes, are used.

[0022] The disadvantage of using a membrane reactor is that reforming and H2 separation must necessarily occur at the same temperature level. Therefore, it is impossible to operate both the reforming and separation processes in the optimum range within a membrane reactor. The interaction is always a compromise based on process requirements: lower temperatures in the membrane reactor have a positive effect on energy efficiency, while higher temperatures have a positive effect on in-situ hydrogen separation. In the case of MeOH reforming, CO2 accumulates in the reaction mixture as H2 is continuously separated during the reforming process. On the other hand, the required reaction heat must be supplied via the heated reactor walls. High CO2 concentrations and high reactor wall temperatures lead to coke deposition, which increases the risk of membrane blockage. To prevent this, additional water must be added to the reaction, which reduces energy efficiency.

[0023] Membrane reactors have attracted considerable academic interest due to their coupling of reaction and H2 separation in process design, but due to the drawbacks mentioned above, they have so far had little practical significance.

[0024] The current consensus among experts is that the energy required for reforming has a negative impact on overall energy efficiency (e.g., Armin Scheuermann, "Flussiger Wasserstoff,Ammoniak oder LOHC-was spricht fur welchen H2-Trager," ["Liquid hydrogen, ammonia or LOHC - Which H2 carrier has the advantage,"] Chemie Technik, May 17, 2022).

[0025] methanol: WO 2004 / 2616 discloses a process consisting of catalytic methanol reforming at 300-500°C and subsequent H separation via pressure swing adsorption (PSA) or with the aid of a palladium alloy membrane. The energy for reforming and hydrogen separation is provided by an internal or external energy source, but the option of using residues and / or extracts of H separation as fuel is not disclosed.

[0026] WO 2003 / 86964 describes a reformer in which methanol reforming and H2 separation from the reformate are carried out using a palladium-based membrane or PSA. For the reforming process, temperatures of 200-700°C are disclosed, and for the methanol reforming process, temperatures of 200-400°C are disclosed. The residue from the H2 separation is burned as an energy source. No information is disclosed regarding the wiring of the necessary heat exchangers. Furthermore, no preheating of the burner air or the methanol is described.

[0027] ammonia: European Patent No. 3,028,990 discloses a method for obtaining a hydrogen / nitrogen mixture by ammonia decomposition. In a first step, liquid ammonia is evaporated; in a second step, gaseous ammonia is split into hydrogen and nitrogen, preferably in a tubular reactor; and in a third step, the product consisting of hydrogen, nitrogen, and unconverted ammonia is cooled in a countercurrent heat exchanger by preheating ambient air, which is then combusted, preferably in a catalytic burner, along with a portion of the product, to provide the heat required for the endothermic reaction of ammonia decomposition. The residual heat of the cooled combustion gases is used to evaporate the liquid ammonia. This method is expected to have an energy efficiency of over 90%. The disadvantage of this method is that the product still contains residual ammonia and is therefore not suitable for fuel cell applications. Therefore, the product is only suitable for purely thermal use, for example in an internal combustion engine. However, this is less efficient than fuel cells.

[0028] British Patent No. 1,079,660 discloses an overall process consisting of catalytic NH3 cleavage and subsequent H2 separation through a Pd alloy membrane. A preferred temperature range of 650-930°C for NH3 cleavage is described. A preferred pressure range is not disclosed. Energy for NH3 vaporization and cleavage is generated electrically.

[0029] The drawback of using electrical energy for NH vaporization and cleavage is that in the best case, this electricity is generated in a downstream fuel cell with an efficiency of up to 70%. This means that not only NH vaporization and cleavage but also H separation is required, and expensive fuel cells not only need to be designed to be larger than if the residue / extract combustion energy were used directly for NH vaporization and cleavage, but also that more NH is required due to the loss of efficiency.

[0030] WO 2018 / 235059 A1 discloses a membrane reactor and method for on-board electricity generation via NH3 cleavage using low-temperature plasma and simultaneous H2 separation using a Pd-Ag membrane. Due to the persistent H2 separation, nearly complete NH3 conversion is achieved even at low temperatures of 200-500 °C and relatively high pressures of 8-10 bar. The cleavage energy is then again electrically supplied.

[0031] WO 02 / 071451 (A2) discloses an H2 generation device for on-board applications. The heart of the system is a compact heat exchanger reactor with many channels. In half of the channels, NH3 is split into N2 and H2 over a ruthenium-nickel catalyst at 550-650°C, while in the other half, fuel is catalytically combusted to provide the heat for NH3 splitting. The reformate from NH3 splitting, consisting primarily of N2 and H2, is converted to electricity in a fuel cell. To protect the fuel cell from unconverted NH3, the process gas first passes through an adsorber bed. The predominantly acidic adsorbent material is not regenerated on-board, but rather replaced. It is proposed that the splitting energy be provided by catalytic combustion of NH3 or, preferably, entrained butane. To initiate the process, the device must be brought to reaction temperature using power from a battery. The disclosed method is suitable for generating electrical power, but is not suitable for generating high purity hydrogen for, for example, refueling stations, due to the lack of separation of N2 and H2.

[0032] L. Lin et al. (L. Lin, Y. Tian, ​​W. Su, Y. Luo, C. Chen, and L. Jiang, "Techno-economic analysis and comprehensive optimization of an on-site hydrogen refueling station system using ammonia: hybrid hydrogen purification with both high H2 purity and high recovery," Sustainable Energy Fuels, vol. 4, pp. 3006-3017, 2020) describe a multi-step process for producing high-purity H2 from NH3 for H2 refueling stations. The results are based on simulations. The paper describes a process that involves catalytic NH3 cleavage at 500 °C, separation of unconverted NH3 in a PSA (pressure swing adsorption) process, separation of the N2 / H2 gas stream by combining the PSA with a membrane process, and compression of the product stream with a purity of 99.97% to a pressure of 900 bar for the refueling station pump. 15.5% of the gas stream from the NH3 cleavage is combusted to provide the required reaction enthalpy. The large number of separation operations and the fact that the reaction enthalpy for NH cleavage is provided by burning the reformate (N, H, and unconverted NH) rather than by burning the extract make the process expensive and result in the need for as little H loss as possible through the residue.

[0033] However, when the entire process chain of evaporation, reforming, and H2 separation is considered in terms of the highest energy efficiency and lowest investment costs, it surprisingly turns out that a combination of reforming, a combination of PSA and TSA for H2 separation, and the best possible heat integration is more effective in terms of the lowest H2 production costs.

[0034] What is needed is a method for producing hydrogen with high purity from methanol or ammonia for hydrogen refueling stations or for distributed supply in industrial applications, in several cost-effective devices and with as little energy loss as possible. Furthermore, it would be advantageous to require fewer expensive materials for catalysts and adsorber. Furthermore, with regard to energy efficiency, it would be advantageous to have as small a difference as possible between the inflow and outflow specific energy flows.

[0035] The present invention relates to a method for obtaining hydrogen from methanol or ammonia, which comprises the steps of: (1) evaporating methanol or ammonia; (2) reforming the methanol or ammonia to form a hydrogen-containing gas mixture; (3) cooling the gaseous product to 25-100°C; (4) separating hydrogen from the cooled gaseous product by a sorption process at a pressure of 1-60 bar and a temperature of 25-100°C; (5) compressing and preheating air in parallel with the first four steps; (6) regenerating the adsorbent with the extract using preheated ambient air; and (7) combusting the extract (tail gas) separated from the adsorbent with air. The combustion gas is passed through at least two different heat exchangers in the direction of the combustion gas flow to (i) first provide reaction heat for reforming the methanol or ammonia, and (ii) subsequently provide evaporation heat for evaporating the feed to the reformer. The reformate preheats ambient air for the regeneration process in a heat exchanger before passing it through the sorption step, the separated hydrogen preheats ambient air for the regeneration process after the sorption process, and / or the combustion gases finally preheat ambient air for the regeneration process as step (iii). [Brief explanation of the drawings]

[0036] [Figure 1] 1 shows an overview of the overall process of the present invention. [Figure 2] 1 shows a schematic representation of the essential components of the method according to the invention. [Figure 3] In the basic case (FIG. 4, variant 1), the symbols used below for streams (S1 to S19), devices (A1 to A9), and heat transfer flows (Q1 to Q6) are shown. [Figure 4] The basic case (variant 1) of the flow pattern of combustion gases is shown. [Figure 5] A second modification is shown. [Figure 6] A third modification is shown. [Figure 7] A fourth modification is shown. [Figure 8] The adsorption process is shown. [Figure 9] The adsorption process is shown. [Figure 10] The second approach is shown. [Figure 11] The LHV (lower heating value) of the feed (MeOH liquid and NH3 liquid) and product (H2 gas) is shown. [Figure 12] 1 shows an example of the method according to the invention for an output of 1000 kg H2 / h according to variant 2, determined on the basis of model calculations. [Figure 13] An example according to the second modification will be shown. [Figure 14] Variation 2 - CG shown [Figure 15] A modified NH3-2- fuel cell is shown. DETAILED DESCRIPTION OF THE INVENTION

[0037] FIG. 1 shows an overview of the overall process of the present invention. FIG. 2 shows a schematic representation of the essential components of the method according to the invention. Figure 3 shows the designations used below for the streams (S1-S19), devices (A1-A9), and heat transfer flows (Q1-Q6) in the base case (Figure 4, Variation 1), which are also listed in Tables 1-4.

[0038] Table 1: Assignment of material flow names used in the text along with the material flow symbols used in the figures.

[0039] [Table 1]

[0040] Table 2: Assignment of equipment names used in the text to equipment symbols used in the figures. Equipment symbols are in the format A1-k, A2-k, etc. and always represent the colder side of the corresponding heat exchanger. Equipment symbols are in the format A1-h, A2-h, etc. and always represent the hotter side of the corresponding heat exchanger.

[0041] [Table 2]

[0042] Table 3: Assignment of heat flow names used in the text to the heat flow symbols used in the figures.

[0043] [Table 3]

[0044] Table 4: Assignment of turbomachinery names used in the text to the symbols used in the figures.

[0045] [Table 4]

[0046] The following description of Figures 3 and 4 refers to the basic case (variant 1).

[0047] First step: The liquid feedstock stream / feed stream S1, which can be an ammonia stream, or a methanol stream, or a mixed stream of methanol and water, is preferably heated in preheater A1 (stream S2) and then vaporized in vaporizer A2 to form reformer feed S3. Heat streams Q1 and Q2 provide these functions. Preheating and vaporization can also be combined in one unit (A1+A2).

[0048] The symbols A1-k and A2-k indicate that these are the cold sides of the heat exchangers A1 and A2, respectively. Accordingly, A1-h and A2-h indicate the warm and hot sides of the heat exchangers A1 and A2, respectively.

[0049] Second step: Reformer feed S3 is advantageously heated in heat exchanger A3 to produce stream S4 which then enters reformer reactor A4. The heat flow Q3 required for heating is advantageously obtained by cooling hot reformate stream S5 to produce cooled reformate stream S6.

[0050] In the reformer reactor, designed as a heat exchanger reactor to introduce the required heat of reaction Q4, the feed components react to form a hot reformate S5 while absorbing energy.

[0051] Third step: Preferably, cooled reformate stream S6 leaving heat exchanger A3 is cooled in heat exchanger A5 to the temperature desired for adsorption of the secondary component. The resulting heat stream Q5 is preferably used to heat air S11 required for desorption and combustion. Cooled reformate stream S7 is fed to adsorption A6-A.

[0052] Fourth step: In the adsorption step A6-A, the secondary components are separated from the hydrogen in the cooled reformate S7 by binding them to a suitable adsorbent, the adsorbent-bound secondary components being referred to hereinafter as the extract.

[0053] Heat is generated during adsorption. Unadsorbed hydrogen leaves the adsorption stage as stream S8, which, due to the resulting heat of adsorption, has a temperature 5-100°C higher than the incoming cooled reformate stream S7. Advantageously, the warm hydrogen stream S8 is cooled in heat exchanger A7 to form cooled H2 product stream S10. The resulting heat stream Q6 is advantageously used to preheat air stream S11.

[0054] Fifth step: Air stream S11 is preferably supplied from the ambient air. The air stream is advantageously warmed in heat exchanger A7 to produce preheated air stream S12, which is then advantageously supplied to compressor A8, which compresses the air sufficiently to overcome all pressure losses before exhaust gases S19 are discharged to the ambient air. Advantageously, compressed air stream S13 is preferably further heated in heat exchanger A5 by heat stream Q5.

[0055] Sixth step: The advantageously compressed and heated air desorbs the extract S9 (shown only in FIG. 3) from the sorbent in desorption stage A6-D. The extract-laden air stream is hereinafter referred to as tail gas stream S15. In this step, the adsorbed absorbent is regenerated.

[0056] Seventh step: Combustible components of tail gas stream S15 are advantageously combusted in burner A9 with the addition of control stream S3a to produce hot combustion gas stream S16. Hot combustion gas stream S16 is then advantageously fed to reformer heat exchanger A4, where it is cooled to warm combustion gas stream S17, primarily to provide the heat of reaction Q4 required for reforming. Warm combustion gas stream S17 is advantageously further cooled in heat exchangers A1 and A2 to heat and vaporize feed stream S1, resulting in heat streams Q1 and Q2. The cooled combustion gases exit the process as exhaust gas stream S19.

[0057] The individual steps are described in detail below.

[0058] First step: methanol: Methanol and optionally water are fed to the evaporator after heating, advantageously with a proportion of water of 0 to 75 mol %, preferably 10 to 70 mol %, particularly preferably 25 to 65 mol %, in particular 40 to 50 mol %, based on the methanol / water mixture, very particularly preferably a molar ratio of methanol to water of 54:46.

[0059] The methanol and / or methanol / water mixture is advantageously evaporated in an evaporator at a pressure of 1 to 60 bar, adjusted in terms of pressure loss and remaining the same throughout the process, to form the gaseous reformer feed. The pressure in the evaporator is advantageously 1 to 30 bar, in particular 2 to 10 bar. The pressure specification provides the skilled person with the temperature required for evaporation.

[0060] ammonia: Alternatively, liquid ammonia is preferably taken from a tank at -35 to 50°C and 1 to 20 bar and, if necessary, brought to a higher pressure by a pump. The liquid ammonia is advantageously converted in an evaporator to gaseous reformer feed at a pressure of 2 to 60 bar, adjusted for pressure loss and remaining the same throughout the process. Advantageously, the pressure in the evaporator is 4 to 40 bar, particularly preferably 6 to 30 bar, in particular 10 to 20 bar. The pressure specification indicates to the skilled person the temperature required for evaporation, advantageously -20 to 100°C.

[0061] As with methanol, the vaporous NH3 stream is advantageously split into a reformer feed that is fed to the reformer and a control stream that is added to the tail gas stream as needed.

[0062] Second step: methanol: The reformer feed, i.e., gaseous methanol or a methanol / water mixture, is then catalytically reformed to gaseous reformate at a temperature of 100-400° C. The methanol reforming temperature is preferably 180-350° C., particularly 240-300° C. A lower methanol reforming temperature increases the H yield while reducing the CO content due to the WGS equilibrium.

[0063] The methanol reformate contains H2, CO, CO2, H2O, and unconverted MeOH or DME. The composition of the gaseous methanol reformate is preferably 55 to 75 mol% H2, 1 to 15 mol% CO, 10 to 25 mol% CO2, 0.5 to 10 mol% H2O, and 0.1 to 20 mol% MeOH and / or DME, and particularly preferably 65 to 75 mol% H2, 6 to 12 mol% CO, 15 to 20 mol% CO2, 1 to 5 mol% H2O, and 0.1 to 5 mol% MeOH and / or DME.

[0064] The reaction yield of the methanol reforming process is advantageously between 70 and 99%, preferably between 80 and 99%, particularly preferably between 85 and 99%.

[0065] During methanol reforming, the reversal of CO2 hydrogenation occurs: 3H2+CO2=CH3OH+H2O DH R 0 = -49 kJ / mol CH3OH The overall reaction follows: CH3OH=2H2+CODH R 0 = +90kJ / mol CH3OH

[0066] According to the present invention, the methanol used may also contain dimethyl ether (C2H6O), typically in a proportion of 1 to 5% by weight, which is simultaneously reformed to methanol in the presence of HO.

[0067] Water reacts with CO according to the following overall reaction: H2O+CO=H2+CO2 DH R0 = -41 kJ / mol CO

[0068] This exothermic reaction is called the water gas shift (WGS) reaction. The water content in methanol can advantageously increase the H yield and reduce the additional energy required for the overall reforming and WGS process.

[0069] The maximum CO produced in the entire process via the WGS reaction and / or combustion of methanol and / or CO corresponds to the CO used in producing methanol from CO and H. Therefore, the process is CO neutral overall.

[0070] Advantageously, no hydrogen stream is extracted during the second step, the reforming process. The second step is therefore advantageously a separate step upstream of the H2 separation. Furthermore, the second step is advantageously separate from and downstream of the first step. An advantageous sequential process step is shown in Figure 3.

[0071] Catalysts for reforming methanol have been described in the prior art (see, for example, F. Gallucci, et al., "Hydrogen Recovery from Methanol Steam Reforming in a Dense Membrane Reactor: Simulation Study," Ind. Eng. Chem. Res. 2004, 43, 2420-2432, and A. Basile, et al., "A dense Pd / Ag membrane reactor for methanol steam reforming: Experimental study," Catalysis Today, 2005, 104, 244-250). For example, a CuO / ZnO / Al2O3 mixture is advantageously used as the active catalyst component, preferably having a composition of 38 wt% CuO, 41 wt% ZnO, and 21 wt% Al2O3, or a mixture having a composition of 31 wt% CuO, 60 wt% ZnO, and 9 wt% Al2O3.

[0072] ammonia: As in the case of methanol, the NH3 vapor stream is advantageously fed to a reformer where it is split into H2 and N2. The energy required for the splitting is advantageously provided by a heat stream. The ammonia reforming is advantageously carried out at a temperature of 100 to 700°C, preferably 200 to 600°C, in particular 300 to 500°C. Advantageously, the ammonia reforming process is carried out at a pressure of 2 to 60 bar, preferably 6 to 30 bar, in particular 10 to 20 bar.

[0073] The gaseous ammonia reformate advantageously contains H2, N2 and unconverted NH3 with the following preferred composition by volume: 60-75% H2, 20-25% N2, 0-20% NH3.

[0074] The yield of the ammonia reforming process is advantageously between 70 and 99%, preferably between 80 and 95%, particularly preferably between 85 and 95%.

[0075] Advantageously, no hydrogen stream is withdrawn during the second step, the reforming process, and therefore the second step is advantageously a separate upstream step. Furthermore, the second step is advantageously separate from and downstream of the first step.

[0076] Catalysts for reforming ammonia have been described in the prior art (see A. Di Carlo, et al., "Ammonia decomposition over commercial Ru / Al2O3 catalyst: An experimental evaluation at different operating pressures and temperatures," International Journal of Hydrogen Energy, 39 (2014), pp. 808-814, or "Ammonia Decomposition on the Process Chain for a Renewable Hydrogen Supply." Chemie Ingenieur Technik, 94 (2022), pp. 1-14). For example, ruthenium is advantageously used as the active catalytic component, as is the ACTA Hypermec 10010 catalyst (Ru / Al2O3). Ni catalysts are particularly active.

[0077] Third step: Both MeOH and ammonia reformate are then cooled in a gas / gas heat exchanger to a preferred temperature of 25-200°C, preferably 35-100°C, especially 40-60°C for H2 separation. Alternatively, ambient air and optionally also the reformer feed gas are advantageously preheated. The optional heat exchanger for preheating the reformer feed is hereinafter referred to as A3, and the heat exchanger for cooling the reformate with simultaneous preheating of ambient air is hereinafter referred to as A5.

[0078] Fourth step: Advantageously, at a temperature of 25 to 200°C, preferably 35 to 100°C, in particular 40 to 60°C, the reformate reaches the adsorber bed for separation of H2.

[0079] For both methanol and ammonia, the H2 concentration in the reformer for H2 separation using a sorbent is advantageously 50-99% by volume, particularly preferably 60-95% by volume, in particular 65-90% by volume.

[0080] In the case of methanol, the CO concentration for H separation using an adsorber is advantageously 0-25 vol%, preferably 0.5-20 vol%, in particular 1-15 vol%. Low CO partial pressures are advantageously achieved by the addition of water, a water-gas shift active catalyst, and / or low temperatures, preferably 150-400°C, in particular 200-250°C.

[0081] If the CO or CO content in the H product gas stream does not meet the requirements of the fuel cell, the H product gas stream can also advantageously be passed through a methanation catalyst bed (see, for example, WO 2004 / 002616 A2).

[0082] In the adsorber, the gaseous product is split into a high-purity H2 product stream, preferably with a purity of greater than 99.99% by volume, and an extract, which is coupled to a solid adsorber (Figures 8 and 9, Adsorption Step). In the case of methanol, the extract consists of CO, CO2, H2O, and unconverted MeOH, as well as small amounts of H2. When ammonia is used, the extract contains not only N2, but also unconverted NH3 and small amounts of H2. The gas space of the sorption device also contains unadsorbed H2.

[0083] Pressure swing adsorption (PSA) and / or temperature swing adsorption (TSA) are preferred adsorption processes, and these separation methods are known to those skilled in the art. In addition, knowledge of suitable adsorbents for separating H from gas mixtures that also contain CO, CO, H0, N, MeOH, or NH is prior art.

[0084] In the adsorption process, a gas mixture is typically introduced at high pressure and low temperature into a fixed-bed reactor packed with adsorbent and flows through the packed bed. One or more components of the mixture (referred to herein as "extract components") are adsorbed. At the bed outlet, the so-called "product component" (in this case, hydrogen) can be removed in concentrated form. After a while, the adsorbent bed becomes largely saturated, and some of the extract component typically leaks out. At this point, the process is switched via a valve so that the outlet for the product component is closed and the outlet for the extract component is opened. This is accompanied by a reduction in pressure and, in the case of temperature swing adsorption, an increase in temperature in the fixed-bed reactor. At low pressure and high temperature, the adsorbed gas can then be desorbed again and recovered at the outlet (see step 6). For example, continuous operation is possible by alternately charging and unloading two fixed-bed reactors packed with adsorbent. It is advantageous to wash the residue of the desorbed extract component from the adsorbent bed with some of the recovered product component to avoid contamination.

[0085] The basics of pressure swing adsorption and suitable adsorbents for separating H from CO-, CO, CH, and HO-containing gas mixtures can be found, for example, in the following publication (Michael Walter; Druckwechseladsorption als Wasserstoffreinigungsverfahren fur Brennstoffzellen-Systeme im kleinen Leistungsbereich [Pressure swing adsorption as a hydrogen purification method for fuel cell systems in the small power range]; Dissertation, University of Duisburg-Essen, 2003). This document also describes the overall system of hydrogen reforming and fuel cells. However, the heat integration detailed therein is based on the use of waste heat by burning exhaust gas from a downstream fuel cell, and not on the combustion of the tail gas resulting from the desorption according to the invention.

[0086] Adsorbents for separating H from N- and NH-containing gas streams are described, for example, in WO 02 / 071451(A2) or L. Lin et al. (L. Lin, Y. Tian, ​​W. Su, Y. Luo, C. Chen and L. Jiang, "Techno-economic analysis and comprehensive optimization of an on-site hydrogen refueling station system using ammonia: hybrid hydrogen purification with both high H purity and high recovery," Sustainable Energy Fuels, Vol. 4, pp. 3006-3017, 2020).

[0087] L. Lin et al. recommend zeolite 13X as an adsorbent for the separation of nitrogen-hydrogen mixtures.

[0088] Carbon-containing adsorbents, such as activated carbon or carbon molecular sieves, and oxide adsorbents, such as zeolites or manganese-magnesium-aluminum oxides, are frequently used. Zeolitic imidazolate structures are also suitable for H2-N2 separation.

[0089] Fifth step: The air is advantageously drawn from the surroundings and compressed to a pressure corresponding to the sum of all pressure losses in the gas lines, starting from the heat exchanger for heating the air, until the air leaves the system as exhaust gas. The sum of all pressure losses can be in the range of 50 mbar to 5 bar. A blower, for example, can be used as the compressor.

[0090] The compressor may also be advantageously located between the desorber and the burner. This has the advantage that a negative pressure is created in the desorber, which promotes desorption. In this arrangement, high temperature levels adversely affect the performance of the compressor.

[0091] To stay below the explosion threshold of 4% hydrogen by volume in air, it may be advantageous to recycle a portion of the oxygen-depleted exhaust gas using a cycle gas compressor prior to the desorption step (see Figure 14, Variant 2-CG).

[0092] Sixth step: The advantageously compressed and heated air, and optionally the circulating gas, desorbs the extract from the sorbent at temperatures of 50 to 700°C, preferably 80 to 600°C, particularly preferably 100 to 500°C, in particular 150 to 400°C, and at pressures of 0.1 to 5 bar, preferably 0.5 to 3 bar, particularly preferably 0.9 to 2 bar (Figures 8 and 9).

[0093] The substantially extractive-free sorbent is then advantageously washed with an inert gas. For the first wash, N2, steam, and / or CO2 can be used. Cycle gas is also suitable as long as the oxygen concentration is low enough not to form explosive mixtures with the hydrogen used for backwashing. After backwashing, the sorption devices contain only substantially extractive-free sorbent and hydrogen. The cooled reformate can now advantageously be recycled to these devices and used for H2 separation.

[0094] During the desorption and scrubbing processes, a total of 11-18% of the product gas (in this case, hydrogen) is typically lost. In general, a larger amount of scrubbing gas means a larger loss of hydrogen (Thomas Joachim Ried, "Further developed CO2 removal by temperature swing adsorption with indirectly tempered adsorbers," Dissertation, Technical University of Munich, August 25, 2020, p. 68).

[0095] However, the method according to the invention makes it possible to achieve lower scrubbing losses of hydrogen in the reformate, in the range of 5-12% by volume.

[0096] Seventh step: The tail gas from the desorption stage using methanol advantageously has the following gas composition: 0.5-5.0 vol.% H, 10-21 vol.% O, 50-80 vol.% N, 0.1-10 mol.% CO, 0.01-10 mol.% CO, 0.01-5 mol.% H0, and 0.001-1 mol.% MeOH, particularly preferably 1-3 vol.% H, 15-20 vol.% O, 70-80 vol.% N, 1-5 mol.% CO, 0.02-5 mol.% CO, 0.1-1 mol.% H0, and 0.001-0.5 mol.% MeOH.

[0097] The tail gas from the desorption stage using ammonia preferably contains the following gas composition: 1-20% by volume H2, 1-10% by volume NH3, 40-95% by volume N2, 5-20% by volume O2, particularly preferably 2-10% by volume H2, 2-6% by volume NH3, 60-85% by volume N2, 12-19% by volume O2.

[0098] The gas mixture of heated air and desorbed extract is fed to a burner which advantageously burns the combustible components in the tail gas, in particular (residual) methanol, carbon monoxide and hydrogen in the case of methanol, and (residual) ammonia and hydrogen in the case of ammonia, with the aid of preheated ambient air for evaporation and reforming, advantageously to cover the energy required for preheating, evaporation and reforming.

[0099] If the heating value of the tail gas is insufficient, a control stream is advantageously withdrawn from the evaporator and fed to the burner.

[0100] The burners may be, for example, atmospheric or catalytic burners. The hot combustion gases, advantageously at a temperature of 500-1300°C in the case of atmospheric burners, and advantageously at a temperature of 300-700°C in the case of catalytic burners, are passed through various heat exchangers to provide (i) heat of reaction for the reforming, (ii) heat of evaporation for the evaporation of the methanol or ammonia, and, if necessary, (iii) preheating of the feedstock.

[0101] The hot combustion gases are advantageously subsequently cooled after leaving the burner. The cooling of the combustion gases is advantageously carried out in such a way that the difference between the two sums of the inlet and outlet specific energy flows, based on the heating value of the hydrogen product flow, is between 0.1 and 10 kWh / kg hydrogen, preferably between 0.2 and 5 kWh / kg hydrogen, more preferably between 0.5 and 4 kWh / kg hydrogen, more preferably between 0.8 and 3 kWh / kg hydrogen, and particularly preferably between 1 and 2 kWh / kg hydrogen.

[0102] In a preferred embodiment, the energy required for evaporation and reforming can be advantageously provided by supplying methanol or ammonia in liquid and gaseous state to the burner in addition to the tail gas. By supplying methanol or ammonia, the entire process can be advantageously started and controlled in stable operating conditions during operation. Mixing with the tail gas can also be carried out before the burner.

[0103] The addition of methanol or ammonia is advantageously controlled via the sensor-detected energy content of the exhaust gas, i.e., the cooled combustion gases exiting the process, and the temperature of the combustion gases from the burner. All of these together provide the energy available for evaporation and reforming. For example, if the burner temperature or exhaust gas volume drops, it is advantageous to add methanol or ammonia to the burner. The amount of methanol or ammonia required can vary widely. The amount of methanol or ammonia fed to the burner is advantageously 0-30%, preferably 0.01-20%, preferably 0.1-15%, and especially 0.5-12% of the amount of methanol or ammonia fed to the entire process.

[0104] In the case of methanol, the hot combustion gases produced in the burner advantageously have a temperature of 350°C to 800°C, preferably 400°C to 700°C, if an atmospheric burner is used, and advantageously have a temperature of 200°C to 600°C, preferably 300°C to 500°C, if a catalytic burner is used.

[0105] In the case of ammonia, the hot combustion gases produced in the burner, if an atmospheric burner is used, advantageously have a temperature of 600°C to 1300°C, preferably 700°C to 1200°C, and if a catalytic burner is used, advantageously have a temperature of 400°C to 700°C, preferably 500°C to 600°C.

[0106] When methanol is used, the combustion gas advantageously contains H2O, CO2, N2, and residual O2. The combustion gas advantageously has the following composition: 10-20% by volume O2, 66-80% by volume N2, 1-10% by volume CO2, 0.1-10% by volume H2O, particularly preferably 15-19% by volume O2, 49-77% by volume N2, 4-9% by volume CO2, 1-9% by volume H2O, in particular 18% by volume O2, 74% by volume N2, 5% by volume CO2, 3% by volume H2O.

[0107] When ammonia is used, the combustion gas advantageously contains N2, O2, and H2O. The combustion gas has the following composition by volume, for example: 78% N2, 9% O2, and 13% H2O.

[0108] In all cases, the composition of the combustion gases is advantageously controlled via the residual O2 concentration. Low O2 values ​​mean a low combustion gas volumetric flow rate (low compression input), but a high initial temperature of the combustion gases. High O2 values ​​(up to 21% by volume) have the opposite effect.

[0109] The flow pattern of the combustion gases is shown in Figure 4 for the base case (variant 1).

[0110] The hot combustion gases are advantageously gradually cooled to near ambient temperature through multiple heat exchangers: (i) reforming, (ii) vaporization of the liquid feedstock, and (iii) optionally preheating of the ammonia, methanol, or methanol-water feed (see Figures 2-6).

[0111] In catalytic burners, the temperature is kept approximately constant throughout the flow path: the temperature on the combustion side is advantageously 1-300°C, preferably 5-50°C, higher than the temperatures in the reformer (200-500°C) and evaporator (130-220°C), i.e., 200-700°C in the reformer and 130-520°C in the evaporator.

[0112] In parallel, the reformate S6, which has a temperature between 200 and 700°C, is advantageously cooled in a heat exchanger A3 by heating the reformer feed S3 (Variant 1, Figure 4). Alternatively, the sensor-detected heat of the reformate S6 can be used to evaporate part of the liquid feed S3 (Variant 3, Figure 6).

[0113] A further advantageous variant is to cool the reformate S6 in the heat exchanger A1 to the desired temperature for adsorption in order to heat the liquid feed S1 and possibly partially evaporate it (variant 2, FIG. 5).

[0114] In variant 1 (FIG. 4), the air drawn into the burner is advantageously preheated. The reformate is thus cooled to a temperature difference with respect to the incoming air flow of 1-200°C, preferably 2-100°C, more preferably 5-80°C, more preferably 8-50°C, in particular 10-40°C. This step is crucial for the energy efficiency of the reformer module.

[0115] In this way, the exhaust gas temperature can be advantageously controlled via the air flow and / or the combustion gas temperature. If the combustion gas temperature is too high, the amount of air taken in is advantageously increased. If the product amount is too low, the control flow S3a is advantageously increased.

[0116] From the viewpoint of high energy efficiency, a small air flow rate is better than a large air flow rate. However, a small air flow rate leads to a high combustion gas temperature, for example, 1100-1200°C. Due to the temperature resistance of materials used for the heat exchanger and gas lines, the combustion gas temperature can be advantageously limited to 1100-1200°C.

[0117] For process control, the exhaust gas volume S19 and H2 production volume S10, as well as the temperatures in gas streams S16, S17, and S18, are preferably measured. The feed flow volume S1 is preferably controlled via the H2 production volume. The gas temperature advantageously regulates the intake air flow S11 and the control flow S3a.

[0118] Downstream fuel cell: A further possibility to further increase the overall efficiency of the whole system is to recycle the off-gas from the downstream fuel cell (which may contain unconverted H2) via desorption to the burner where it is used for energy production (modified NH3-2-fuel cell, see Figure 14).

[0119] Therefore, the present invention also provides a method for generating electricity from methanol or ammonia, comprising the steps of: evaporating methanol or ammonia in a first step; reforming it into a hydrogen-containing gas mixture in a second step; cooling the gaseous product to 25-200°C in a third step; separating hydrogen, together with nitrogen in the case of ammonia, from the cooled gaseous product in a fourth step by a sorption process at a pressure of 1-60 bar and a temperature of 25-200°C; generating electricity from the separated hydrogen in a fuel cell in a fifth step in parallel with the first step. in a sixth step, the adsorbent on which the extract is adsorbed is regenerated with off-gas from the anode side and / or the cathode side of the fuel cell; in a seventh step, the extract separated from the adsorbent is combusted with air, wherein the combustion gas is passed in the flow direction of the combustion gas through at least three different heat exchangers to (i) first provide reaction heat for reforming methanol or ammonia, (ii) then provide evaporation heat for evaporating the reformer feed, and finally (iii) preheat the extract separated from the air and / or the adsorbent.

[0120] Preferably, the combustion gas passes through at least four different heat exchangers in the direction of flow of the combustion gas: (i) first to provide heat of reaction for reforming the methanol or ammonia, then (ii) to provide heat of evaporation for evaporating the reformer feed, then (iii) to preheat the extract separated from the adsorbent, and finally (iv) to preheat the air.

[0121] Preferably, in the sixth step, the adsorbent that has adsorbed the extract is regenerated with off-gases from the anode and cathode sides of the fuel cell.

[0122] Preferably, the reformate preheats the feed before the evaporator and / or the reformer feed after the evaporator in a heat exchanger before entering the sorption step. Particularly preferably, the reformate first heats the reformer feed and then heats the feed before entering the sorption step.

[0123] Preferably, the air required for the burner and fuel cell is supplied via a single compressor.

[0124] Ammonia is preferably used as the feed. Suitable adsorbent materials for ammonia contain salt complexes of 3d transition metals. These materials are preferably manganese, iron, cobalt, nickel, copper, or zinc. Other suitable materials are impregnated activated carbons such as AddsorbTMVB1, which is a phosphoric acid-impregnated activated carbon, or zeolites such as Fe / HBEA, for example, with a silicon to aluminum ratio of 12.5.

[0125] advantage Adsorption processes such as PSA and TSA have the disadvantage that the tail gas produced during the regeneration of the adsorbent containing the extract contains approximately 11-18% of the hydrogen produced (Thomas Joachim Ried, "Further developed CO removal by temperature swing adsorption with indirectly tempered adsorbers", Dissertation, Technical University of Munich, August 25, 2020, p. 68), and therefore represents a loss. This disadvantage does not occur in the method according to the invention, since the calorific value of the hydrogen in the tail gas can be used to heat the heat-consuming reforming process.

[0126] Since no additional energy is supplied from the outside, and no excess energy is released to the outside, apart from electrical consumers such as air compressors, the H2 product stream must have the same calorific value as the liquid feedstocks, methanol or ammonia, in the theoretical limit. Therefore, in the process according to the invention, theoretically no conversion energy is lost. Losses occur only due to heat dissipation to the surroundings through the equipment walls, as the discharge stream is warmer than the feed stream, and due to mechanical power losses in the air conveying elements.

[0127] Analysis parameters: In the process according to the invention, the external energy balance is determined solely by the energy stored in the inlet and outlet streams (Figure 10). In a theoretically ideal case where the methanol / water or ammonia feed streams, and the air and electricity for air compression, introduce the same power into the process as the H2 product outlet stream and exhaust gas export, this would result in 100% energy efficiency for the process.

[0128] Energy efficiency may be determined using two approaches, with only the second approach providing instructions for achieving low energy efficiency.

[0129] First approach: In the first approach, the chemical energy content of the feed stream and the H2 product stream is used to determine the energy efficiency, which can be determined from their respective heating values ​​(LHV = lower heating value) multiplied by the corresponding mass flow rates.

[0130] Chemical energy content of the feed stream E F、LHV is the mass flow rate m F and calorific value LHV F It is calculated from E F、LHV =m F * LHV F

[0131] Energy content E of the H2 product stream H2 is therefore the mass flow rate m H2 and calorific value LHV H2 It is calculated from E H2、LHV =m H2 * LHV H2

[0132] Chemical conversion loss ΔE V is E F、LHV and E H2、LHV It is the difference between ΔE V、chem =E F、LHV -E H2、LHV

[0133] For this approach, the feed (MeOH 液体 and NH3 液体 ) and the product (H2 ガス ) is required. These are known from the literature and can be determined as shown in Figure 11. In addition, the mass flow rate of the input energy carrier m F , and the mass flow rate of the extracted energy carriers H2m H2 must be known.

[0134] Chemical energy efficiency η of the conversion of MeOH or NH3 to H2 EN、chem can be defined as follows: η EN、chem =E H2、LHV / E F、LHV =(E F、LHV -ΔE V、chem ) / E F、LHV =1-ΔE V、chem / E F、LHV

[0135] Energy conversion loss ΔE V、tot Regarding the gross power P1 input to the method, グロス needs to be considered additionally. ΔE V、tot = ΔE V、chem +P1 グロス

[0136] Energy efficiency η of MeOH or NH3 conversion to H2 EN、tot can be defined as follows: η EN、tot =E H2、LHV / (E F、LHV +ΣP グロス )=(E F、LHV -ΔE V、tot ) / E F、LHV =1-ΔE V、tot / E F、LHV

[0137] Second approach: In the second approach, the energy flow into and out of the process is balanced across the balance boundaries of the method to determine the energy efficiency (FIG. 10).

[0138] The energy flow rate in the calculation example is calculated as follows: Gaseous flow: E ガス =m ガス * cp ガス、平均 * (TT 0 ) Liquid flow:E 液体 =m 液体 * cp ガス、平均 * (TT 0 )-ΔH V where cp ガス、平均 is T 0 ~TT 0 is the average specific heat of the stream in the temperature range of TT 0 is the reference temperature, which is fixed at 25°C in the following calculations. The ambient temperature at which the streams enter and leave the balance region of the method is also fixed at 25°C. This means that these TT 0 = 0, there is no need to consider the flow rate of the inflowing gas. 0 Only exiting gas streams with a temperature T higher than TT are considered. 0 = 0 also applies, so the enthalpy of vaporization of water ΔH V、H2OOnly the evaporation of MeOH and NH3 needs to be considered. MeOH、液体 and LHV NH3、液体 (See Figure 11) The enthalpy of vaporization ΔH V、MeOH and ΔH V、NH3 Therefore, the respective heat values ​​LHV 液体 Each energy-carrying feed stream can be treated like a gas stream and therefore has a value of 0.

[0139] In the second approach, the conversion loss ΔE V is calculated as follows (see FIG. 10): ΔE V = Energy flow out - Energy flow in ΔE V =E H2 +E EG +ΣQ V -m H2O * (-ΔH V、H2O )-P1 net E H2 =m H2 * cp H2、平均 * (T H2 -T 0 ) E EG =m EG * cp EG、平均 * (T EG -T 0 )

[0140] ΔE V To miniaturize, good thermal integration is required, which means a small E H2 and small E EG, As reflected in the above, the flow rate and pressure loss of the combustion gases must be low, the air conveying elements must be highly efficient, and the device must have good thermal insulation.

[0141] Regarding the electrical energy for compressor P1, the gross power Pグロス Of which, net electricity P ネット Only the net power output can be taken into account because only the net power output is reflected as an energy component in the product stream. V =P1 グロス -P1 ネット cannot be used in the surrounding area.

[0142] However, the overall efficiency of the method, η EN、tot Regarding electrical energy, gross power P1 グロス The total energy conversion loss ΔE V、tot then, ΔE V、tot = ΔE V +P1 グロス

[0143] This is the overall efficiency of the method, η EN、tot Result: η EN、tot =1-ΔE V、tot / E F、LHV and specific total conversion loss ΔE V、spec is the mass flow rate of H2 product m H2 Related to: ΔE V、spec = ΔE V、tot / m H2

[0144] Small energy conversion loss ΔE V、tot It can be seen that not only should the temperature difference between the outflow and the ambient be kept small, but also the corresponding mass flow rate. At a constant H2 product flow, only the exhaust gas flow is affected. In addition, the inflow water flow should be kept small, so that as much heat flow as possible is used in the method, and the device should be able to reduce the heat loss flow rate ΣQ V In addition, high mechanical efficiency in air flow machines reduces the energy conversion loss ΔE V、tot Reduce.

[0145] Based on this knowledge, the objective is to develop a process for converting MeOH and NH3 to hydrogen that has the lowest possible enthalpy to the effluent and does not generate heat loss flows.

[0146] Specifically, this means connecting the flows with respect to their heat exchange such that the difference between the sum of the two inlet and outlet specific energy flows is between 1 and 5 kWh / kg hydrogen, based on the heating value of the hydrogen product flow.

[0147] The method according to the invention achieves this objective. The method according to the invention advantageously has an overall energy efficiency η of 80-99%, preferably 90-98%. EN、tot The method according to the invention allows for a total energy conversion loss ΔE V、tot is 2 to 8 kWh / kg H2, preferably 2.5 to 4 kg / kg H2, and when ammonia is used, it is 0.5 to 3 kWh / kg H2, preferably 1.0 to 2 kg / kg H2.

[0148] Table 5 again lists the terms in tabular form.

[0149] Table 5: Assignment of energy flow names used in the text along with symbols used in the figures.

[0150] [Table 5]

[0151] First example - methanol: FIG. 12 shows an example of the method according to the invention for an output of 1000 kg H2 / h according to variant 2, determined on the basis of model calculations.

[0152] This example is the result of a thermodynamic simulation in terms of quantity and energy using a BASF internal simulator similar to the simulation program Aspen Plus.

[0153] This example is calculated without heat loss through the walls of the method apparatus.

[0154] According to the process according to the invention, 6505 kg of methanol and 3723 kg of water must be fed to the process per hour at a temperature of 25° C. and a pressure of 3 bar.

[0155] Countercurrent flow of the liquid feed and warm reformate in the second reformate cooler preheats the liquid feed mixture and cools the reformate to an adsorption temperature of 40°C. This requires a heat transfer capacity of 912 kW. The liquid feed mixture has a boiling point of 116°C at 3 bar. 10,227 kg of untreated condensed steam is fed to the reformer as reformer feed and 1 kg / hr is fed to the burner as a control flow.

[0156] In the reformer, the raw condensed steam is brought to a reaction temperature of 250°C and catalytically reformed to 70.8 vol% H, 3.8 vol% CO, and 21.0 vol% CO. The equilibrium MeOH conversion at 250°C and 3 bar is theoretically 99.9%. The reformate additionally contains 4.3 vol% unconverted H0 and 0.1 vol% unconverted MeOH. The reforming process requires 3750 kW of thermal energy.

[0157] The reformate is then cooled to 167° C. in the first reformate cooler. In turn, the vaporous raw condensate is heated from 116° C. to 230° C. Heating requires a thermal power of 608 kW.

[0158] In the adsorber, 9227 kg / hr of extract is adsorbed at 3 bar and 40°C and / or stored in the gas void volume of the adsorber particles. This example is calculated with a 14% H loss, as described in Syed Naqvi, "Hydrogen Production," PEP Report 32C, SRI Consulting, September 2007. This means that 163 kg / hr of hydrogen is lost as scrubbing losses and 1000 kg / hr of hydrogen leaves the adsorber as the product stream. Heat is released during adsorption. The warm product stream has a temperature of 59°C and is cooled to 40°C in the product cooler. This requires a heat transfer capacity of 75 kW.

[0159] The extract stream from the adsorber contains 25.3 vol.% H2, 9.8 vol.% CO, and 53.8 vol.% CO2, 11.0 vol.% H2O, and 0.1 vol.% MeOH.

[0160] In the desorption step, the extract stream is entrained by a heated air stream and then, after compression, combusted in a burner.

[0161] This requires 100,564 kg / hr of air, which is first heated to 28° C. in a product cooler or air heater and then to 134° C. in an exhaust gas cooler before entering the desorber to take up the extract. The extract-laden air leaves the desorber as tail gas in an amount of 109,791 kg / hr with a composition of 2.1 vol.% H, 19.2 vol.% O, 72.5 vol.% N, 0.8 vol.% CO, and 4.5 vol.% CO, 0.9 vol.% HO, and 0.01 vol.% MeOH.

[0162] The desorption process requires heat, so the temperature in the desorber drops to 122°C.

[0163] To overcome all stream losses, the tail gas is compressed to 1.2 bar in the air conveying element. Calculations are based on a mechanical efficiency of 90% and an isentropic efficiency of 75%. Compression requires an electrical power of 962 kW. During compression, the temperature of the tail gas increases from 122°C to 149°C.

[0164] In the burner, the compressed tail gas is combusted with a controlled flow of 1 kg / hr from the evaporator, producing hot combustion gases with a temperature of 395°C. This combustion gas is fed to the reformer, where it provides 3750 kW of heat output for the reforming reaction and heats the gaseous reformer feed to 230°C to 250°C.

[0165] The warm combustion gases from the reformer have a temperature of 281°C and are cooled to 154°C in the evaporator, then further cooled in countercurrent with supplied ambient air to an exhaust gas temperature of 58°C.

[0166] Energy efficiency of the entire process η En、tot Therefore, η EN、tot =1-ΔE V、tot / E F =1-(ΔE V、chem +ΣΔP V / E F =1-[(36053-33320)+962] / 36053=0.898.

[0167] Chemical conversion loss ΔE V、chem is calculated as follows: ΔE V、chem =E V、H2 +E V、AG +ΣQ V -(-m H2O * ΔH V、H2O )-ΣP ネット =60+1018+0+2521-0.9 * 962=2733kW ΔE V、tot = ΔE V、chem +ΣP グロス =2733+962=3695kW

[0168] From this the energy efficiency of the entire process can be calculated. η EN、tot =1-ΔE V、tot / E F =1-3695 / 36053=0.898

[0169] H2 production material flow rate m H2 and its associated specific total conversion loss ΔE V、spec can then be expressed as: ΔE V、spec = ΔE V、tot / m H2 =3695kWh / kg H2

[0170] The total energy efficiency, excluding heat loss through the equipment walls, is η EN、tot =90.0%.

[0171] Second Example - Ammonia An example of the second modification is shown in FIG.

[0172] This example is the result of a thermodynamic simulation in terms of quantity and energy using a BASF internal simulator similar to the simulation program Aspen Plus.

[0173] This example is calculated without heat loss through the walls of the method apparatus.

[0174] According to the process according to the invention, 7110 kg of liquid ammonia must be fed to the process per hour at a temperature of 25° C. and a pressure of 10 bar.

[0175] Countercurrent flow of liquid ammonia and warm reformate in the second reformate cooler preheats the liquid ammonia to 92°C and cools the reformate to the adsorption temperature of 40°C. This requires a heat transfer capacity of 680 kW. Liquid ammonia has a boiling point of 25°C at 10 bar. 7109 kg of ammonia vapor is supplied to the reformer as reformer feed and 1 kg / hr is supplied to the burner as a control flow.

[0176] In the reformer, ammonia vapor is brought to a reaction temperature of 400°C and catalytically reformed to 71.9% H2 by volume and 24.0% N2 by volume. The equilibrium NH3 conversion at 400°C and 10 bar is theoretically 92%. The reformate also contains 4.1% unconverted NH3 by volume. The reforming requires 5681 kW of thermal energy.

[0177] The reformate is then cooled to 144° C. in the first reformate cooler. In turn, the vaporous raw condensate is heated from 25° C. to 380° C. Heating requires a thermal power of 1702 kW.

[0178] In the adsorber, 6109 kg / hr of extract is adsorbed at 10 bar and 40°C or stored in the gas void volume of the adsorber particles. This example is calculated with a 14% H loss, as described in Syed Naqvi, "Hydrogen Production," PEP Report 32C, SRI Consulting, September 2007. This means that 163 kg / hr of hydrogen is lost as scrubbing losses and 1000 kg / hr of hydrogen leaves the adsorber as the product stream. Heat is released during adsorption. The warm product stream has a temperature of 88°C and is cooled to 40°C in the product cooler. This requires a heat transfer capacity of 191 kW.

[0179] The extract stream from the adsorber contains 26.4 vol.% H2, 62.9 vol.% N2, and 10.7 vol.% NH3.

[0180] In the desorption step, the extract stream is entrained by a heated air stream and then, after compression, combusted in a burner.

[0181] This requires 21,715 kg / hr of air, which is first heated to 56° C. in a product cooler or air heater and then to 131° C. in an exhaust gas cooler before entering the desorber to take up the extract. The extract-laden air leaves the desorber as tail gas in an amount of 27,824 kg / hr with a composition of 7.6 vol.% H, 14.9 vol.% O, 74.4 vol.% N, 3.1 vol.% NH.

[0182] The desorption process requires heat, so the temperature in the desorber drops to 82°C.

[0183] To overcome all stream losses, the tail gas is compressed to 1.2 bar in the air conveying element. Calculations are based on a mechanical efficiency of 90% and an isentropic efficiency of 75%. Compression requires an electrical power of 241 kW. During compression, the temperature of the tail gas increases from 82°C to 107°C.

[0184] In the burner, the compressed tail gas is combusted with a controlled flow of 1 kg / hr from the evaporator, producing hot combustion gases with a temperature of 1000°C. This combustion gas is fed to the reformer, where it provides 5681 kW of heat output for the reforming reaction and heats the gaseous reformer feed to 380°C to 400°C.

[0185] The warm combustion gases from the reformer have a temperature of 410°C and are cooled to 274°C in the evaporator, then further cooled in countercurrent with supplied ambient air to an exhaust gas temperature of 222°C.

[0186] Energy efficiency of the entire process η En、tot Therefore, η EN、tot =1-ΔE V、tot / E F =1-(ΔE V、chem +ΣΔP V ) / EF =1-[(34853-33320)+241] / 34853=0.949.

[0187] Chemical conversion loss ΔE V、chem is calculated as follows: ΔE V、chem =E V、H2 +E V、AG +ΣQ V -(-m H2O * ΔH V、H2O )-ΣP ネット =60+1684+0+0-0.9 * 241=1527kW ΔE V、tot = ΔE V、chem +ΣP グロス =1527+241=1768kW

[0188] From this the energy efficiency of the entire process can be calculated. η EN、tot =1-ΔE V、tot / E F =1-1768 / 34853=0.949

[0189] H2 production material flow rate m H2 and its associated specific total conversion loss ΔE V、spec can then be expressed as: ΔE V、spec = ΔE V、tot / m H2 =1768kWh / kg H2

[0190] The total energy efficiency, excluding heat loss through the equipment walls, is η EN、tot =94.9%.

[0191] Variations: There are several options for heat integration, all of which have in common that the combustion gases from the burner pass in a first step through the heat exchanger of the reformer (A4) and in a second step through the heat exchanger of the evaporator (A2).

[0192] In Variant 1 (Figure 4), the vaporized feedstock is heated in a first reformate cooler before entering the reformer, and the warm reformate is cooled in a second reformate cooler by heating preheated air. Residual cooling of the combustion gases is achieved by preheating the feed stream.

[0193] In Variant 2 (FIG. 5), similar to Variant 1, the vaporized feedstock is heated in a first reformate cooler before entering the reformer. However, the warm reformate is not cooled by heated air, but rather by the incoming liquid feedstock. Residual cooling of the cooled combustion gases after the vaporizer is achieved by heating the heated air.

[0194] In variant 3 (Figure 6), the evaporator functions as the first reformate cooler, the air is heated in a second reformate cooler before entering the desorber, and the warm combustion gases are cooled to exhaust gas temperatures in a feed preheater before exiting the process.

[0195] In Variant 4 (FIG. 7), as in Variant 3, the evaporator functions as the first reformate cooler, and the air is heated in a second reformate cooler before entering the desorber. As in Variant 2, the warm reformate is cooled by the incoming liquid feedstock. Residual cooling of the cooled combustion gases after the evaporator is achieved by heating the heated air.

[0196] Table 7 compares the results for the individual variants.

[0197] The symbols should be read as follows: MeOH-2, for example, means variant 2 when using methanol; variant NH3-1 means variant 1 when using ammonia; and NH3-2-00 refers to variant 2 with ammonia as the energy carrier when the temperature difference between the hot and cold streams in the heat exchanger is zero. However, this requires an infinitely large heat exchanger surface. Hence, an overall efficiency η of 98.1%. en、totThis variant with represents a theoretical limit.

[0198] If the H2 loss through the tail gas is higher, the gas mixture in the desorber may be explosive. In this case, it is advisable to use a cycle gas compressor to recirculate part of the exhaust gas to the desorber (variant NH3-2-CG, Figure 14).

[0199] Overall efficiency of H2 loss through tail gas η en、tot The effect of H2 loss on the overall efficiency was also investigated. In the base case, a 14% H2 loss is assumed. Cases were also calculated for 5%, 8%, and 11%. Reducing H2 loss from 14% to 11% still results in a significant increase in overall efficiency from 94.9% to 97.1%, but further reductions do not improve overall efficiency very much.

[0200] An important result of these model calculations is therefore the use indication to reduce H losses via tail gas by approximately 11%, which is possible with the method according to the invention, since here desorption is not carried out by applying negative pressure as described in the prior art, but rather the extracted gas stream (cycle gas or combustion air) can be used to reduce the partial pressure of the extracted substances in the gas phase (tail gas).

[0201] The combination of the method according to the invention with a fuel cell (FC) for power generation brings advantages for power generation, which are illustrated using the case of ammonia (variant NH3-2-fuel cell, see Figure 15). If the fuel cell is operated with a mixture of H2 and N2 instead of pure H2, the driving potential for power generation is reduced at the same pressure. However, this can be compensated for by a 25% pressure increase. In this way, the associated lower NH3 conversion rate in the reformer is inconsequential, as the unconverted NH3 is converted in the burner to heat, which is necessary for the reforming and evaporation operations. The separation of NH3 from H2-containing gas requires much less effort than the separation of N2 and H2, so the adsorption effort is significantly reduced. The reason is that the boiling point difference between NH3 and H2 is much greater than that between N2 and H2. In addition, the alkaline effect of NH3 can be very well utilized. ·The difficult separation of H2 and N2 occurs naturally in fuel cells through a membrane that is only permeable to H+ ions. The exhaust gas from a fuel cell is particularly suitable for desorption because it contains almost no oxygen. In this case, no recycle gas is required to prevent explosive concentrations. Instead of two air compressors, one for the method according to the invention and one for the fuel cell, only one is needed. A larger air compressor is always more cost-effective than two smaller ones. In addition, the larger the compressor size, the greater the efficiency. The use of the exhaust gas stream from the fuel cell as a cleaning gas during desorption also improves efficiency. At 95.2%, it is higher than the efficiency of variants with the same heat integration. In comparison, the variant NH3-2 achieves an efficiency of 94.9% and the recycle gas variant NH3-2-CG achieves an efficiency of 94.8%.

[0202] Table 7: Comparison of variants for a 1 kg / hr H2 product flow. Enthalpy of H2 product flow H2 = 33.32 kW

[0203] [Table 6]

[0204] Comparison of NH3-2 variant with prior art: A comparison of the energy efficiency of the method according to the invention with the prior art provides an overview of the energetic and therefore economic advantages of the present invention. British Patent No. 1,079,660 65% WO 2018 / 235059(A1) <78% L. Lin et al. <80% WO 02 / 071451(A2) 85% Lamb et al. 90% EP 3,028,990 >90% (The product in this case is only a gas mixture of H2 / N2 / NH3) This invention 94-97%

[0205] The method according to the invention for producing high purity hydrogen has the highest energy efficiency.

Claims

1. 1. A method for obtaining hydrogen from methanol or ammonia, comprising: a first step of evaporating methanol or ammonia; a second step of reforming the methanol or ammonia into a hydrogen-containing gas mixture; a third step of cooling the gaseous product to 25-100°C; a fourth step of separating the hydrogen from the cooled gaseous product by a sorption process at a pressure of 1-60 bar and a temperature of 25-100°C; a fifth step of compressing air in parallel with the first four steps and preheating it together with the hydrogen separated after the sorption process; a sixth step of regenerating an adsorbent with the extract using the preheated air; and a seventh step of combusting the extract separated from the adsorbent together with the air, wherein combustion gases are passed through at least two different heat exchangers in the direction of flow of the combustion gases to (i) first provide heat of reaction for reforming the methanol or ammonia, and (ii) subsequently provide heat of evaporation for evaporating the feed of a reformer.

2. 2. A method according to claim 1, characterized in that the seventh step, i.e. desorption, is carried out at a temperature between 80 and 500° C. and a pressure between 0.5 and 3 bar.

3. 3. The method according to claim 1 or 2, characterized in that the reforming step is carried out at 180 to 350°C for methanol and at 300 to 500°C for ammonia.

4. 4. The method according to claim 1, wherein the reformate further heats the preheated air in a heat exchanger before entering the sorption step.

5. The method according to any one of claims 1 to 3, characterized in that the combustion gas then further heats the preheated air as step (iii).

6. 4. The method according to any one of claims 1 to 3, characterized in that the reformate is preheated in a heat exchanger before the feed enters the sorption step.

7. 7. The method of claim 6, wherein the combustion gas is passed through at least three different heat exchangers in the flow direction of the combustion gas to (i) first provide the reaction heat for reforming the methanol or ammonia, (ii) then provide the evaporation heat for evaporating the reformer feed, and (iii) finally preheat the air for the regeneration.

8. 4. The method according to claim 1, wherein the combustion gas is passed through at least three different heat exchangers in a flow direction of the combustion gas to (i) first provide the heat of reaction for reforming the methanol or ammonia, (ii) then provide the heat of evaporation for evaporating the feed of the reformer, and (iii) finally preheat the feed upstream of an evaporator.

9. 9. The method of claim 8, wherein the reformate provides additional heat of vaporization for vaporizing the reformer feed in a heat exchanger before entering the sorption step.

10. 4. The method according to any one of claims 1 to 3, characterized in that the reformate provides additional heat of evaporation for evaporating the reformer feed in a heat exchanger before entering the sorption step, and preheats the feed in a further heat exchanger before the evaporator.

11. 11. The method of claim 10, wherein the combustion gas is passed through at least three different heat exchangers in a flow direction of the combustion gas to (i) first provide the heat of reaction for reforming the methanol or ammonia, (ii) then provide the heat of evaporation for evaporating the reformer feed, and (iii) finally further heat the preheated air.

12. A method for producing electricity from methanol or ammonia, comprising the steps of: in a first step, evaporating the methanol or ammonia; in a second step, reforming it into a hydrogen-containing gas mixture; in a third step, cooling the gaseous product to 25-200°C; in a fourth step, separating the hydrogen, together with nitrogen in the case of ammonia, from the cooled gaseous product by a sorption process at a pressure of 1-60 bar and a temperature of 25-200°C; and producing electricity from the separated hydrogen in a fuel cell; and in a fifth step, compressing and pre-treating air in parallel with the first step. and in a sixth step, regenerating the adsorbent with the extract adsorbed thereon with off-gas from the anode side and / or the cathode side of the fuel cell. In a seventh step, combusting the extract separated from the adsorbent with the air, wherein the combustion gas is passed through at least three different heat exchangers in the flow direction of the combustion gas to (i) first provide reaction heat for reforming the methanol or ammonia, (ii) then provide evaporation heat for evaporating a reformer feed, and finally (iii) preheat the air and / or the extract separated from the adsorbent.

13. 13. The process according to claim 12, characterized in that ammonia is used as the feed.

14. 14. The method according to claim 12 or 13, wherein in the sixth step, the adsorbent having adsorbed the extract is regenerated with the off-gas on the anode side and the cathode side of the fuel cell.

15. 15. The method according to any one of claims 12 to 14, characterized in that the reformate preheats the feed before the evaporator and / or the feed of the reformer after the evaporator in a heat exchanger before entering the sorption step.

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