Method and device for obtaining high-purity hydrogen from methanol or ammonia

Abstract

The invention relates to a method for obtaining hydrogen from methanol or ammonia. The invention is characterized in that in a first step, methanol or ammonia is evaporated; in a second step, the methanol or ammonia is reformed in order to form a hydrogen-containing gas mixture; in a third step, the gaseous reformate is cooled to 25 to 100 °C; in a fourth step, the hydrogen is separated from the cooled gaseous reformate by means of a sorption process at a pressure of 1 to 60 bar and at a temperature of 25 to 100 °C; in a fifth step, air is compressed and preheated in parallel with the first four steps; in a sixth step, the adsorbent loaded with the extract is regenerated using the pre-heated ambient air; and in a seventh step, the extract separated from the adsorbent, the tail gas, is combusted using the air. The combustion gases are guided via at least two different heat exchangers, in order to (i) first provide the reaction heat for reforming the methanol or the ammonia and (ii) subsequently provide the evaporation heat for evaporating the reformer feed in the flow direction of the combustion gases. The reformate pre-heats the ambient air for the regeneration process in a heat exchanger prior to entering the sorption step, the separated hydrogen pre-heats the ambient air for the regeneration process after the sorption process, and / or the combustion gases finally pre-heat the ambient air for the regeneration process as step (iii).

Description

[0001] Process and device for producing high-purity hydrogen from methanol or ammonia

[0002] Description

[0003] The present invention relates to a process for the production of hydrogen from methanol or ammonia, which is characterized in that methanol or ammonia is evaporated in a first step and reformed in a second step to a hydrogen-containing gas mixture, in a third step the gaseous reformate is cooled to 25 to 100°C, in a fourth step the hydrogen is separated from the cooled gaseous reformate at a pressure of 1 to 60 bar and a temperature of 25 to 100°C by a sorption process, wherein parallel to the first four steps in a fifth step air is compressed and preheated, in a sixth step the adsorbent loaded with the extract is regenerated with the preheated ambient air, and in a seventh step the extract separated from the adsorbent, the tail gas, is combusted with the air, wherein the combustion gases are passed through at least two different heat exchangers,to provide, in the flow direction of the fuel gases, (i) first the reaction heat for the reforming of the methanol or ammonia and (ii) subsequently the heat of vaporization for the vaporization of the reformer feed, whereby the separated hydrogen preheats the ambient air for regeneration after the sorption process.

[0004] Hydrogen offers the desired prerequisites to become a key factor in the energy supply of the future. The transport sector and industry, in particular, face the major challenge of becoming more climate-friendly. According to calculations by the Federal Network Agency, approximately 4.3 million tonnes of hydrogen will be needed annually in Germany alone by 2045 to achieve climate neutrality. This corresponds to a thermal output of approximately 144 TWh per year (approval of the scenario framework 2023-2037 / 2045, Federal Network Agency, July 2022, (https: / / www.netzausbau.de / Wissen / Ausbaubedarf / Szenariorahmen / de.html). By comparison, today's large-scale plants have a maximum annual production capacity of less than 0.1 million tonnes of hydrogen. Furthermore, the feedstock is almost exclusively fossil.

[0005] Hydrogen will be produced using renewable energies in the future. Since the meteorological 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 used as a raw condensate, a mixture of methanol and reaction water.

[0006] What is needed is a technology that can recover hydrogen from methanol or ammonia with high efficiency and at low cost.

[0007] In order to use hydrogen in industry as a chemical raw material or in the transport sector in fuel cell applications, the hydrogen must be of very high quality, as impurities affect catalysts and membranes.

[0008] Hydrogen is currently predominantly produced centrally in steam methane reforming (SMR) production units. If the production site and the utilization site are located far apart, hydrogen must be highly compressed (up to 350 bar) and, in rare cases, liquefied, in order to transport it to the location where it is needed, for example, to a hydrogen filling station, using appropriate transport vehicles. However, transporting hydrogen by vehicle is uneconomical and environmentally unfriendly, as larger hydrogen filling stations would have to be supplied daily by a truck.

[0009] In parallel with vehicle transport, there are a few dedicated hydrogen pipelines. However, to supply hydrogen to filling stations on a large scale, a separate, dense hydrogen pipeline network similar to the natural gas network would have to be built. Such pipeline networks, however, have very high infrastructure costs and require complex approval procedures, making their implementation in the near future unlikely.

[0010] Water electrolysis requires a very high amount of electricity, which, due to the poor storage capacity of hydrogen at filling stations and in industry, must be met on an as-needed basis using the available grid power. Since grid power demand will almost double in the future if, in addition to today's electricity consumers, the car sector and heat generation are to be powered entirely by electricity, not only wind power and photovoltaic capacities will need to be significantly expanded, but the power grid as well.

[0011] To minimize the large amount of land and sea space required for this, along with the negative environmental impacts (this applies especially to wind turbines), it makes sense to relocate hydrogen production to regions where the meteorological conditions are significantly better and the open spaces are considerably larger than in Germany. Efforts are already underway to produce this renewable energy demand in countries with very favorable conditions, such as the MENA (Middle East and North Africa) states. One example of this is the NEOM HELIOS project in Saudi Arabia, currently the world's largest green hydrogen / ammonia project.

[0012] Hydrogen production costs are significantly lower in windy and sunny countries than in Germany. To take advantage of this cost advantage, transportation costs must be low. The two hydrogen carriers methanol and, above all, ammonia meet this requirement. Liquid organic hydrogen carriers (LOHCs) also meet this requirement, albeit with some limitations.

[0013] Methanol (MeOH) is a large-scale basic chemical and an excellent energy carrier due to its high energy density of 19.9 MJ / kg. Unlike hydrogen, methanol can be transported cost-effectively (O. Machhammer, "Renewable Electricity from Germany or e-Fuels from Chile: What Should Future Mobility Be Based on?" Chemie Ingenieur Technik, No. 4, 2021). The existing crude oil transport infrastructure can be used for transportation.

[0014] Methanol is currently still primarily recycled as a basic chemical, for example, to produce formaldehyde, acetic acid, methyl chloride, methyl methacrylate, and methylamines. The energy balance plays a minor role in these processes; the added value of the resulting products is crucial.

[0015] Ammonia (NH3) is a basic chemical produced on a large scale, for example, for the production of fertilizer. Ammonia is a good energy carrier; at 18.6 MJ / kg, it has approximately the same mass-related energy density as methanol (MeOH) at 19.9 MJ / kg. Ammonia has a boiling point of -33°C and can be transported at ambient temperature in 10-bar low-pressure containers.

[0016] A key feature of future energy sources will be their low carbon footprint. In the case of NH3, in addition to renewably produced hydrogen (H2), nitrogen (N2) is also required, which is highly concentrated in the atmosphere (approximately 80%) and can therefore be easily extracted via an air separation plant.

[0017] In today's primary use of ammonia for material purposes, such as fertilizer, the energy balance plays a subordinate role. The effect of the fertilizer itself is crucial.

[0018] Known processes for the separation of 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, as well as membrane processes.

[0019] The hydrogen can be made available at a filling station, for example, for refueling fuel cell vehicles. For this purpose, the hydrogen is compressed to the required pressure of 950 bar for intermediate storage and cooled to the required temperature of -40°C during refueling.

[0020] However, if methanol or ammonia is used as an energy source, the energy balance of the entire process plays a key role. The entire process, from the reforming of the methanol or ammonia to the release of the hydrogen, should advantageously exhibit low energy losses in order to retain as much of the originally used energy as possible.

[0021] The operation of fuel cells (FCs) requires hydrogen of very high purity (>99.99%). The production of hydrogen of the highest purity from methanol or ammonia requires several process steps: the evaporation and splitting of methanol or ammonia, and the separation of the high-purity hydrogen from the resulting gas mixture. The thermal energy required for the evaporation and splitting must either be supplied externally or provided by burning a portion of the methanol, ammonia, or a portion of the reforming products.

[0022] The state of the art focuses primarily on maximum conversion in reforming and optimized hydrogen separation.

[0023] To maximize conversion, membrane reactors are currently predominantly proposed, and for optimized hydrogen separation, membrane processes or a combination of PSA (pressure swing adsorption) and membrane processes are used.

[0024] A disadvantage of using membrane reactors is that reforming and H2 separation must necessarily take place at the same temperature level. Therefore, it is not possible to operate both the reforming and separation processes individually within their optimal range. The interaction is always a process-related compromise: A lower temperature in the membrane reactor has a positive effect on energy efficiency, while a higher temperature has a positive effect on in-situ hydrogen separation. Because H2 is continuously separated during the reforming process in MeOH reforming, CO2 accumulates in the reaction mixture. The necessary reaction heat must also be supplied via the heated reactor walls. A high CO2 concentration and hot reactor walls lead to coke deposits. This increases the risk of membrane blockage.To prevent this, additional water must be added to the reaction, which reduces the energetic efficiency.

[0025] Membrane reactors are of academic interest due to the process-technical coupling of reaction and H2 separation; however, due to the disadvantages mentioned above, they have so far had little practical significance.

[0026] Furthermore, the current common opinion among experts is that the energy required for reforming negatively impacts the overall energy efficiency (e.g. Armin Scheuermann, “Liquid hydrogen, ammonia or LOHC - what speaks for which H2 carrier”, Chemie Technik, May 17, 2022).

[0027] Methanol:

[0028] WO 2004 / 2616 discloses a process consisting of catalytic methanol reforming at 300 to 500°C followed by hydrogen separation via pressure swing adsorption (PSA) or with the aid of palladium alloy membranes. The energy for the reforming and hydrogen separation is provided by an internal or external energy source, although the option of using the retentate or extract of the hydrogen separation as fuel is not disclosed.

[0029] WO 2003 / 86964 describes a reforming device in which methanol reforming and H2 separation from the reformate are carried out using a palladium-based membrane or a PSA. Temperatures of 200 to 700°C are disclosed for reforming, and 200 to 400°C for methanol reforming. The retentate from the H2 separation is combusted as the energy source. No details are disclosed regarding the wiring of the required heat exchangers. Furthermore, no preheating of the burner air or the methanol is described.

[0030] Ammonia:

[0031] EP 3,028,990 discloses a process for producing a hydrogen-nitrogen mixture by ammonia cracking. In a first step, liquid ammonia is evaporated. In a second step, the gaseous ammonia is cracked into hydrogen and nitrogen, preferably in a tubular reactor. In a third step, the reformate consisting of hydrogen, nitrogen, and unreacted ammonia is cooled in a countercurrent heat exchanger by preheating the ambient air. This air is then combusted with a portion of the reformate, preferably in a catalytic burner, thus providing the necessary heat for the endothermic ammonia cracking reaction. The residual heat in the cooled fuel gas is used to evaporate the liquid ammonia. Energy efficiency rates of greater than 90% are expected for this process.The disadvantage of this process is that the reformate still contains residual ammonia and is therefore unsuitable for fuel cell applications. The reformate is therefore only suitable for purely thermal use, e.g., in an internal combustion engine. However, this has a lower efficiency than a fuel cell. GB 1,079,660 discloses an overall process consisting of catalytic NH3 splitting and subsequent H2 separation via Pd-alloy membranes. A preferred temperature range of 650 to 930°C for the NH3 splitting is described; preferred pressure ranges are not disclosed. The energy for the NH3 evaporation and splitting is generated electrically.

[0032] The disadvantage of using electrical energy for NH3 evaporation and splitting is that, in the best case scenario, this electricity is generated in the downstream fuel cell with an efficiency of a maximum of 70%. This not only requires larger NH3 evaporation and splitting, as well as H2 separation and the expensive fuel cell to be designed larger than in the case of directly using the combustion energy of the retentate / extract for NH3 evaporation and splitting, but also requires more NH3 due to the loss in efficiency.

[0033] WO 2018 / 235059 A1 discloses a membrane reactor and a method for generating electricity on-board via NH3 splitting using low-temperature plasma and simultaneous H2 separation with Pd-Ag membranes. Due to the continuous H2 separation, almost complete NH3 conversion is achieved even at low temperatures of 200 to 500°C and at relatively high pressures of 8 to 10 bar. The splitting energy is again supplied electrically.

[0034] WO 02 / 071451 A2 discloses an H2-generating apparatus for on-board applications. Its core is a compact heat exchanger reactor with multiple channels. While NH3 is split into N2 and H2 at 550 to 650 °C over ruthenium-nickel catalysts in one half of the channels, a fuel is catalytically combusted in the other half to provide the heat for the NH3 splitting. The reformate from the 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 is first passed over an adsorber bed. The preferably acidic adsorber material is not regenerated on-board, but rather exchanged. It is proposed that the splitting energy be provided by catalytic combustion of NH3 or, preferably, by catalytic combustion of entrained butane.To start the process, the device is heated to reaction temperature using electricity from a battery. The disclosed method is suitable for generating electrical power, but not for producing high-purity hydrogen, for example, for use at gas stations, because it lacks the separation of N2 and H2.

[0035] 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-stage process for producing high-purity H2 from NH3 for an H2 refueling station. The results are based on simulations. The process involves catalytic NH3 decomposition at 500°C, separation of the unreacted NH3 in a PSA (pressure swing adsorption), separation of the N2 / H2 gas stream using a combination of PSA and membrane processes, and compression of the product stream with a purity of 99.97% to a pressure of 900 bar for the gas station pump. 15.5% of the gas stream from the NH3 decomposition is combusted to cover the required reaction enthalpy.The large number of separation operations and the fact that the reaction enthalpy for NH3 splitting is provided by burning the reformate (N2, H2, and unreacted NH3) rather than by burning the extract makes the process expensive and necessitates the loss of as little H2 as possible via the retentate. However, if the entire process chain of evaporation, reforming, and H2 separation is considered from the perspective of maximum energy efficiency and lowest investment costs, it surprisingly turns out that a combination of reforming and a combination of PSA and TSA for H2 separation, as well as the most optimal heat integration, is more effective in terms of lowest production costs.

[0036] Therefore, a process is needed to produce high-purity hydrogen from methanol or ammonia for hydrogen refueling stations or for the decentralized supply of industrial applications. This process produces hydrogen in a small number of cost-effective devices and with as little energy loss as possible. Furthermore, the low requirement for expensive materials for the catalysts and adsorbers is advantageous. Furthermore, it is advantageous for energy efficiency if the specific energy flows in and out differ as little as possible.

[0037] The present invention relates to a process for the production of hydrogen from methanol or ammonia, which is characterized in that methanol or ammonia is evaporated in a first step, reformed in a second step to a hydrogen-containing gas mixture, in a third step the gaseous reformate is cooled to 25 to 100°C and in a fourth step the hydrogen is separated from the cooled gaseous reformate at a pressure of 1 to 60 bar and a temperature of 25 to 100°C by a sorption process, wherein in parallel to the first four steps, in a fifth step air is compressed and preheated, in a sixth step the adsorbent loaded with the extract is regenerated with the preheated ambient air, and in a seventh step the extract separated from the adsorbent, the tail gas, is combusted with the air, wherein the combustion gases are passed through at least two different heat exchangers,in order to provide, in the flow direction of the fuel gases, (i) first the reaction heat for the reforming of the methanol or ammonia and (ii) subsequently the heat of vaporization for vaporizing the reformer feed, wherein the reformate preheats the ambient air for regeneration in a heat exchanger before entering the sorption step, the separated hydrogen preheats the ambient air for regeneration after the sorption process and / or the fuel gases finally preheat the ambient air for regeneration as step (iii).

[0038] Figure 1 shows an overview of the overall process of the invention.

[0039] Figure 2 shows schematically the essential components of the method according to the invention.

[0040] Figure 3 shows the designations used for the base case (Figure 4, Variant 1) for the streams (S1 to S19), the equipment (A1 to A9), and the heat transfer streams (Q1 to Q6). These are also listed in Tables 1 to 4.

[0041] Table 1 : Assignment of the material flow names used in the text with the material flow designations used in the figures.

[0042] Table 2: Correlation of the device names used in the text with the device designations used in the figures. Device designations in the form A1-k, A2-k, etc. always represent the flow side of the colder stream in the corresponding heat exchanger. Device designations in the form A1-h, A2-h, etc. always represent the flow side of the hotter stream in the corresponding heat exchanger.

[0043] Table 3: Correspondence of the heat flow names used in the text with the heat flow designations used in the figures.

[0044] Table 4: Correspondence of the names used in the text for turbomachines with the designations used in the figures.

[0045] The following explanation of Figures 3 and 4 refers to the base case (variant 1).

[0046] First step:

[0047] The liquid feed stream S1, which can be an ammonia stream, a methanol stream, or a mixture of methanol and water, is preferably heated in a preheater A1 (stream S2) and then evaporated in an evaporator A2 to form the reformer feed S3. Heat flows Q1 and Q2 serve this purpose. Preheating and evaporation can also be combined in a single device (A1 + A2).

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

[0049] Second step:

[0050] Reformer feed S3 is advantageously heated in heat exchanger A3 to form stream S4 before entering reformer reactor A4. The heat flow Q3 required for heating is advantageously obtained by cooling the hot reformate stream S5 to form the cooled reformate stream S6.

[0051] In the reformer reactor, which is designed as a heat exchanger reactor to introduce the necessary reaction heat Q4, the feed components react to form the hot reformate S5 while absorbing energy.

[0052] Third step:

[0053] Preferably, the cooled reformate stream S6 exiting heat exchanger A3 is cooled in a heat exchanger A5 to the desired temperature for adsorption of the secondary components. The resulting heat flow Q5 is preferably used to heat the air S11 required for desorption and combustion. The cooled reformate stream S7 is fed to adsorption A6-A.

[0054] Fourth step:

[0055] In the adsorption stage A6-A, the secondary components are separated from the hydrogen in the cooled reformate S7 by binding them to a suitable adsorbent. The secondary components bound to the adsorbent are referred to below as the extract.

[0056] Heat is generated during adsorption. The unadsorbed hydrogen leaves the adsorption stage as stream S8, which, due to the resulting adsorption heat, has a temperature 5 to 100°C higher than the incoming cooled reformate stream S7. The warm hydrogen stream S8 is advantageously cooled in heat exchanger A7 to form the cooled H2 product stream S10. The resulting heat stream Q6 is advantageously used to preheat the air stream S11.

[0057] Fifth step:

[0058] The air stream S11 is preferably fed from the ambient air. After the air stream has been advantageously warmed in heat exchanger A7 to form the preheated air stream S12, it is advantageously fed to a compressor A8. This compressor compresses the air to such an extent that all pressure losses can be overcome until the exhaust gas S19 is released to the ambient air. The advantageously compressed air stream S13 is further heated preferably in heat exchanger A5 by the heat stream Q5. Sixth step:

[0059] The advantageously compressed and heated air desorbs the extract S9 (shown only in Figure 3) from the sorbent in the desorption stage A6-D. The air stream laden with the extract is referred to below as tail gas stream S15. In this step, the laden sorbent is regenerated.

[0060] Seventh step:

[0061] The combustible components of tail gas stream S15 are advantageously combusted in a burner A9 with the addition of control stream S3a to form the hot fuel gas stream S16. The hot fuel gas stream S16 is then advantageously fed to the reformer heat exchanger A4, where the hot fuel gas stream S16 cools to form the warm fuel gas stream S17, primarily to cover the reaction heat Q4 required for reforming. The warm fuel gas stream S17 is advantageously further cooled in heat exchangers A1 and A2 to heat and vaporize the feed stream S1 using the resulting heat flows Q1 and Q2. The cooled fuel gas leaves the process as offgas stream S19.

[0062] The individual steps are explained in more detail below.

[0063] First step:

[0064] Methanol:

[0065] Methanol and optionally water are fed to an evaporator after heating. The water content is advantageously between 0 and 75 mol% based on the methanol-water mixture, preferably between 10 and 70 mol%, particularly preferably between 25 and 65 mol%, especially between 40 and 50 mol%, and very particularly preferably between 54 and 46 mol%.

[0066] The methanol or the methanol-water mixture is advantageously evaporated in an evaporator at pressures between 1 and 60 bar, which are constant throughout the process after adjusting for pressure loss, to form the gaseous reformer feed. The pressure in the evaporator is advantageously between 1 and 30 bar, especially between 2 and 10 bar. The temperatures required for evaporation can be determined by the person skilled in the art from the pressure specifications.

[0067] Ammonia:

[0068] Alternatively, liquid ammonia is taken from a tank, advantageously at -35 to 50°C and 1 to 20 bar, and brought to higher pressures with the aid of a pump if necessary. The liquid ammonia is advantageously converted into the gaseous reformer feed in the evaporator at pressures between 2 and 60 bar, which are constant throughout the process, adjusted for pressure loss. The pressure in the evaporator is advantageously between 4 and 40 bar, particularly preferably between 6 and 30 bar, and especially between 10 and 20 bar. The pressure data provide the skilled person with the necessary temperatures for evaporation, advantageously between -20°C and 100°C.

[0069] As in the case of methanol, the vaporous NH3 stream is advantageously split into a reformer feed that is fed to the reformer and a control stream that is mixed into the tail gas stream as required.

[0070] Second step:

[0071] Methanol:

[0072] The reformer feed, i.e., the gaseous methanol or methanol-water mixture, is then catalytically reformed at temperatures between 100 and 400°C to form a gaseous reformate. The methanol reforming temperature is preferably between 180 and 350°C, especially between 240 and 300°C. Low methanol reforming temperatures increase the H2 yield at the expense of the CO content due to the WGS equilibrium.

[0073] The methanol reformate comprises H2, CO, CO2, H2O, and unreacted MeOH or DME. The composition of the gaseous methanol reformate preferably consists of 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, 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.

[0074] The conversion of the methanol reforming is advantageously 70 to 99%, preferably 80 to 99%, particularly preferably 85 to 99%.

[0075] During the reforming of methanol, the reversal of CO2 hydrogenation takes place

[0076] 3 H2 + CO2 = CH3OH + H2O DHR° = -49 kJ / mol CH3OH according to the following overall reaction equation instead of

[0077] CH3OH = 2 H2 + CO DH R ° = +90 kJ / mol CH3OH

[0078] According to the invention, the methanol used may also contain proportions of dimethyl ether (C2H6O), typically 1 to 5 wt.%. Dimethyl ether is simultaneously reformed to methanol in the presence of H2O.

[0079] Water reacts with CO according to the following overall reaction equation:

[0080] H2O + CO = H2 + CO2 DH R ° = -41 kJ / mol CO

[0081] This exothermic reaction is called the water-gas shift (WGS) reaction. The water content in methanol can advantageously increase the hydrogen yield and reduce the additional energy required for the overall reforming and WGS process.

[0082] The maximum CO2 produced in the overall process via the WGS reaction and / or combustion of methanol and / or CO corresponds to the CO2 used in methanol production from CO2 and H2. The overall process is therefore CO2 neutral.

[0083] Advantageously, no hydrogen stream is withdrawn during the second step, reforming. Thus, the second step is advantageously a separate step, preceding the H2 separation. Furthermore, the second step is advantageously separate from the first step and downstream of it. The advantageous successive process steps are shown in Figure 3.

[0084] Catalysts for reforming methanol are 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 / g membrane reactor for methanol steam reforming: Experimental study," Catalysis Today, 2005, 104, 244-250). For example, CuO / ZnO / Al2O3 mixtures are used as active catalyst components; advantageously in the composition of 38 wt% CuO, 41 wt% ZnO, and 21 wt% Al2O3, or mixtures in the composition of 31 wt% CuO, 60 wt% ZnO, and 9 wt% Al2O3. Ammonia:

[0085] Analogous to 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 covered by a heat flow. Ammonia reforming advantageously takes place at temperatures of 100 to 700°C, preferably 200 to 600°C, in particular between 300 and 500°C. Ammonia reforming advantageously takes place at a pressure of 2 to 60 bar, preferably 6 to 30 bar, in particular 10 to 20 bar.

[0086] The gaseous ammonia reformate advantageously contains H2, N2 and unreacted NH3 in the following preferred composition: 60 to 75 vol% H2, 20 to 25 vol% N2, 0 to 20 vol% NH3.

[0087] The conversion of the ammonia reforming is advantageously 70 to 99%, preferably 80 to 95%, particularly preferably 85 to 95%.

[0088] Advantageously, no hydrogen stream is withdrawn during the second step, reforming. Thus, the second step is advantageously a separate upstream step. Furthermore, the second step is advantageously separate from the first step and downstream of it.

[0089] Catalysts for reforming ammonia are described in the state of the art (see A. Di Carlo, et.al ., “Ammonia decomposition over commercial Ru / AI2O3 catalyst: An experimental evaluation at different operative 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 used as an active catalyst component; ACTA Hypermec 10010 catalyst_(Ru / AI2O3) is advantageous. Ni catalysts are particularly active.

[0090] Third step:

[0091] Both the MeOH and ammonia reformate are then cooled in a gas / gas heat exchanger to the preferred temperature of 25 to 200°C, preferably 35 to 100°C, in particular 40 to 60°C, for H2 separation. In return, the ambient air and optionally also the reformer feed gas are advantageously preheated. The optional heat exchanger for preheating the reformer feed is referred to below as A3, and the heat exchanger for cooling the reformate with advantageous simultaneous preheating of the ambient air is referred to below as A5.

[0092] Fourth step:

[0093] At a temperature of advantageously 25 to 200°C, preferably 35 to 100°C, in particular 40 to 60°C, the reformate reaches the adsorber bed for the separation of H2.

[0094] In the case of both methanol and ammonia, the H2 concentration in the reformate for H2 separation with sorbents is advantageously between 50 and 99 vol%, particularly preferably between 60 and 95 vol% and in particular between 65 and 90 vol%.

[0095] In the case of methanol, the CO concentration for H2 separation with adsorbers is advantageously between 0 and 25 vol%, preferably between 0.5 and 20 vol%, and in particular between 1 and 15 vol%. A low CO partial pressure is advantageously achieved by the addition of water, a water-gas shift catalyst, and / or low temperatures, preferably 150 to 400°C, in particular 200 to 250°C. In the event that the CO or CO2 content in the H2 product gas stream does not meet the requirements of the fuel cell, the H2 product gas stream can also advantageously be passed over a methanation catalyst bed (see, for example, WO 2004 / 002616 A2).

[0096] In the adsorber, the gaseous reformate is split into a high-purity H2 product stream, preferably with a purity of > 99.99 vol% H2, and the extract, which is bound to the solid adsorber (Figures 8 and 9, adsorption step). In the case of methanol, the extract consists of CO, CO2, H2O, and unreacted MeOH, as well as small amounts of H2. When ammonia is used, the extract contains not only N2 but also unreacted NH3 and small amounts of H2. The gas space of the sorption apparatus also contains unadsorbed H2.

[0097] Pressure swing adsorption (PSA) and / or temperature swing adsorption (TSA) are preferred adsorption processes; these separation processes are known to those skilled in the art. Furthermore, knowledge of suitable adsorbents for separating H2 from gas mixtures that also contain CO, CO2, H2O, N2, MeOH, or NH3 is state of the art.

[0098] During the adsorption step, a gas mixture is typically introduced into a fixed-bed reactor filled with the adsorbent at elevated pressure and low temperatures. It then flows through the fixed-bed reactor. One or more components of the mixture (referred to here 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 adsorber bed is largely saturated, and some of the extract components typically escape. At this point, the process is switched via valves so that the outlet for the product component is closed and an outlet for the extract components is opened. This is accompanied by a reduction in pressure and - in the case of temperature swing adsorption - supported by an increase in temperature in the fixed-bed reactor.At the low pressure and high temperature, the adsorbed gas is then desorbed again and can be recovered at the outlet (see step six). For example, two alternately loaded and unloaded fixed-bed reactors filled with adsorbent enable continuous operation. To expel the residual desorbed extract component from the adsorber bed, it is advantageous to flush with a portion of the recovered product component to avoid contamination.

[0099] The fundamentals of pressure swing adsorption and suitable adsorbents for the separation of H2 from CO, CO2, CH4, and H2O-containing gas mixtures can be found, for example, in the following reference (Michael Walter; Pressure Swing Adsorption as a Hydrogen Purification Process for Fuel Cell Systems in the Low Power Range; Dissertation, University of Duisburg-Essen, 2003). This document also describes the entire system consisting of hydrogen reforming and a fuel cell. However, the heat integration described therein is based on the utilization of waste heat by combusting the exhaust gases from the downstream fuel cell and not on the inventive combustion of the tail gas resulting from desorption.

[0100] Adsorbents for the separation of H2 from N2- and NH3-containing guest streams are listed, for example, in WO 02 / 071451 A2 or in L. Lin et al. (L. Lin, Y. Tian, ​​W. Su, Y. Luo, C. Chen and L. Jiang, “Techno-economic analysis and comprehensive optimizaition of an on-site hydrogen refuelling station system unsing ammonia: hybrid hydrogen purification with both high H2 purity and high recovery,” Sustainable Energy Fuels, Vol. 4, pp. 3006-3017, 2020).

[0101] L. Lin et al. recommend Zeolite 13 X as an adsorbent for the separation of nitrogen-hydrogen mixtures.

[0102] Carbonaceous 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, for example, are also suitable for H2-N2 separation.

[0103] Fifth step:

[0104] Air is advantageously drawn in from the ambient air and compressed to a pressure that corresponds to the sum of all pressure losses in the gas line, from the heat exchanger used to heat the air to its exit from the system as exhaust gas. The sum of all pressure losses can range from 50 mbar to 5 bar. Air blowers, for example, can be used as compressors.

[0105] The compressor can also be advantageously located between the desorber and the burner. This has the advantage of creating a negative pressure in the desorber, which promotes desorption. The higher temperature level in this arrangement would be detrimental to the compressor performance.

[0106] It may be advantageous to recirculate part of the oxygen-depleted exhaust gas using a cycle gas compressor before the desorption stage (see Figure 14, variant 2-KG) in order to keep the explosion limit for hydrogen below 4 vol% in air.

[0107] Sixth step:

[0108] The advantageously compressed and heated air, and optionally the cycle gas, desorb the extract from the sorbent at temperatures between 50 and 700 °C, preferably at 80 to 600 °C and particularly preferably at 100 to 500 °C, in particular between 150 and 400 °C, and at pressures between 0.1 and 5 bar, preferably at 0.5 and 3 bar, particularly preferably at 0.9 to 2 bar (Figures 8 and 9).

[0109] The largely extract-free sorbent is then advantageously flushed with an inert gas. For the initial flush, N2, steam, and / or CO2 can be used, for example. Recycle gas is also suitable, as long as the oxygen concentration is low enough to prevent an explosive mixture from forming with the hydrogen used for backflushing. After backflushing, the sorption apparatus contains only largely extract-free sorbent and hydrogen. The cooled reformate can then advantageously be fed back into these for H2 separation.

[0110] During desorption and purging processes, a total of 11 to 18% of the product gas (in this case, hydrogen) is typically lost. In general, larger purge gas volumes result in greater hydrogen losses (Thomas Joachim Ried, "Advanced CO2 Removal by Temperature Swing Adsorption with Indirectly Temperature-Controlled Adsorbers," dissertation, Technical University of Munich, August 25, 2020, p. 68).

[0111] However, the process according to the invention can make it possible to achieve lower purge losses in the range of 5 to 12 vol.% of the hydrogen in the reformate

[0112] Seventh step: The tail gas from the desorption stage using methanol advantageously has the following gas composition: 0.5 to 5.0 vol% H2, 10 to 21 vol% O2, 50 to 80 vol% N2, 0.1 to 10 mol% CO, 0.01 to 10 mol% CO2, 0.01 to 5 mol% H2O and 0.001 to 1 mol% MeOH, particularly preferably 1 to 3 vol% H2, 15 to 20 vol% O2, 70 to 80 vol% N2, 1 to 5 mol% CO, 0.02 to 5 mol% CO2, 0.1 to 1 mol% H2O and 0.001 to 0.5 mol% MeOH.

[0113] The tail gas from the desorption stage preferably contains the following gas composition using ammonia: 1 to 20 vol% H2, 1 to 10 vol% NH3, 40 to 95 vol% N2, 5 to 20 vol% O2, particularly preferably 2 to 10 vol% H2, 2 to 6 vol% NH3, 60 to 85 vol% N2, 12 to 19 vol% O2,

[0114] The gas mixture of heated air and desorbed extract is fed to a burner, which advantageously combusts 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 the preheated ambient air in order to cover the required energy for evaporation and reforming, advantageous for preheating, evaporation and reforming.

[0115] If the calorific value of the tail gas is insufficient, a control stream is advantageously drawn from the evaporator and fed to the burner.

[0116] The burner can be, for example, an atmospheric burner or a catalytic burner. The hot fuel gas, advantageously at a temperature of 500 to 1300°C in the case of an atmospheric burner and advantageously at a temperature of 300 to 700°C in the case of a catalytic burner, is passed through various heat exchangers to provide (i) the reaction heat for reforming, (ii) the heat of vaporization for vaporizing the methanol or ammonia, and, if necessary, (iii) the preheating of the feedstock.

[0117] The hot fuel gas is advantageously cooled successively after leaving the burner. The cooling of the fuel gas advantageously occurs such that the difference between the two sums of the incoming and outgoing specific energy flows, based on the calorific value of the hydrogen product stream, 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.

[0118] In a preferred embodiment, the energy required for evaporation and reforming can be provided by advantageously feeding methanol or ammonia in liquid and gaseous states to the burner in addition to the tail gas. By feeding methanol or ammonia, the entire process can be advantageously started up and controlled in a stable operating state during operation. The admixture can also advantageously take place upstream of the burner into the tail gas.

[0119] The addition of methanol or ammonia is advantageously controlled via the sensible energy content of the exhaust gas, i.e. the cooled fuel gas leaving the process, and the temperature of the fuel gases from the burner. All of this together results in the energy available for evaporation and reforming. If, for example, the burner temperature or the amount of exhaust gas drops, then methanol or ammonia will advantageously be fed to the burner. The required amount of methanol or ammonia can vary greatly. The amount of methanol or ammonia fed to the burner is advantageously between 0 and 30%, preferably between 0.01 and 20%, preferably between 0.1 and 15%, in particular between 0.5 and 12% of the amount of methanol or ammonia fed to the overall process.

[0120] The hot fuel gas produced in the burner advantageously has a temperature of 350°C to 800°C, preferably 400°C to 700°C, when using an atmospheric burner for the methanol case and advantageously has a temperature of 200 to 600°C, preferably 300 to 500°C, when using a catalytic burner.

[0121] The hot fuel gas produced in the burner advantageously has a temperature of 600°C to 1300°C, preferably 700°C to 1200°C, when using an atmospheric burner for the ammonia case, and advantageously has a temperature of 400 to 700°C, preferably 500 to 600°C, when using a catalytic burner.

[0122] When using methanol, the fuel gas advantageously contains H2O, CO2, N2 and residual O2. The fuel gas advantageously has the following composition: 10 to 20 vol% O2, 66 to 80 vol% N2, 1 to 10 vol% CO2, 0.1 to 10 vol% H2O, particularly preferably 15 to 19 vol% O2, 49 to 77 vol% N2, 4 to 9 vol% CO2, 1 to 9 vol% H2O, in particular 18 vol% O2, 74 vol% N2, 5 vol% CO2, 3 vol% H2O.

[0123] The fuel gas advantageously contains N2, O2, and H2O when using ammonia. The fuel gas has, for example, the following composition: 78 vol% N2, 9 vol% O2, and 13 vol% H2O.

[0124] In all cases, the composition of the fuel gas is advantageously controlled by the residual O2 concentration. Low O2 values ​​result in low fuel gas volume flows (low compression effort) but a high initial fuel gas temperature. High O2 values ​​(maximum 21 vol%) have the opposite effect.

[0125] The flow pattern of the fuel gas is shown in Figure 4 for the base case (variant 1).

[0126] The hot fuel gas advantageously passes successively through several heat exchangers, (i) the reforming (ii), the evaporation of the liquid feedstock and (iii) optionally the pre-heating of the ammonia, methanol or methanol-water feed and is gradually cooled to near ambient temperature (see Figures 2 to 6).

[0127] In the catalytic burner, the temperature remains approximately constant along the entire flow path. The temperature on the combustion side is advantageously 1 to 300°C, preferably 5 to 50°C, higher than the temperature in the reformer (200 to 500°C) and the evaporator (130 to 220°C). This means that the temperature on the combustion side is 200 to 700°C in the reformer and 130 to 520°C in the evaporator.

[0128] In parallel, the reformate S6, which has a temperature of 200 to 700°C, is advantageously cooled in the heat exchanger A3 by heating the reformer feed S3 (variant 1, Figure 4). Alternatively, the sensible heat of the reformate S6 can be used to evaporate part of the liquid feed S3 (variant 3, Figure 6).

[0129] Another advantageous variant is to cool the reformate S6 in heat exchanger A1 to the desired temperature for adsorption in order to heat the liquid feed S1 and possibly even partially evaporate it (variant 2, Figure 5). In variant 1 (Figure 4), the air drawn in for the burner is advantageously preheated. The reformate is thus cooled to a temperature difference from the incoming air stream of 1 to 200°C, preferably 2 to 100°C, more preferably 5 to 80°C, even more preferably 8 to 50°C, in particular 10 to 40°C. This step is of great importance for the energy efficiency of the reformer module.

[0130] In the process, the exhaust gas temperatures can be advantageously controlled via the air flow rate and / or the combustion gas temperatures. If the combustion gas temperature is too high, the intake air flow is advantageously increased. If the product flow rate is too low, the control flow S3a is advantageously increased.

[0131] In terms of high energy efficiency, small airflows are better than large ones. However, small airflows result in high combustion gas temperatures, e.g., 1100 to 1200°C. Due to the temperature resistance of the materials used for the heat exchangers and gas lines, the combustion gas temperature should be limited to 1100 to 1200°C.

[0132] To control the process, the exhaust gas flow rate S19 and the H2 product flow rate S10, as well as the temperatures in the gas streams S16, S17, and S18, are preferably measured. The inlet flow rate S1 is preferably controlled via the H2 product flow rate. The gas temperatures advantageously control the intake air flow rate S11 and the control flow rate S3a.

[0133] Downstream fuel cell:

[0134] Another possibility to further increase the overall efficiency of the entire system is to recirculate the offgas from a downstream fuel cell, which may contain unconverted H2, into the burner via desorption in order to use it there for energy production (see variant NH3-2-BSZ, Figure 14).

[0135] The present invention thus also includes a process for producing electricity from methanol or ammonia, which is characterized in that methanol or ammonia is evaporated in a first step, reformed in a second step to a hydrogen-containing gas mixture, in a third step the gaseous reformate is cooled to 25 to 200°C, in a fourth step the hydrogen and, in the case of ammonia, together with the nitrogen, are separated from the cooled gaseous reformate at a pressure of 1 to 60 bar and a temperature of 25 to 200°C by a sorption process, and further electricity is produced from the separated hydrogen in a fuel cell, wherein, in parallel to the first steps, air is compressed and preheated in a fifth step, in a sixth step the adsorbent loaded with the extract is regenerated with the off-gases from the anode and / or cathode side of the fuel cell,and in a seventh step, the extract separated from the adsorbent is combusted with the air, wherein the fuel gases are passed through at least three different heat exchangers in order to provide, in the flow direction of the fuel gases, (i) first the reaction heat for the reforming of the methanol or ammonia and (ii) subsequently the heat of vaporization for vaporizing the reformer feed, and (iii) finally to preheat the air and / or the extract separated from the adsorbent.

[0136] Preferably, the fuel gases are passed through at least four different heat exchangers in order to (i) first provide the reaction heat for reforming the methanol or ammonia, (ii) then the heat of vaporization for vaporizing the reformer feed, (iii) then preheat the extract separated from the adsorbent, and (iv) finally preheat the air. Preferably, in the sixth step, the adsorbent loaded with the extract is regenerated with the off-gases from the anode and cathode sides of the fuel cell.

[0137] 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 the feed before entering the sorption step.

[0138] Preferably, the air required for the burner and the fuel cell is supplied via a single compressor.

[0139] Ammonia is preferably used as the feed. Suitable adsorbent materials for ammonia contain a complex salt of the 3d transition metals. These materials are preferably manganese, iron, cobalt, nickel, copper, or zinc. Other suitable materials are impregnated activated carbons such as AddsorbTMVBI, a phosphoric acid-impregnated activated carbon, or zeolites such as Fe / HBEA, for example, with a silicon to aluminum ratio of 12.5.

[0140] Advantages:

[0141] Adsorption processes such as PSA and TSA have the disadvantage that the tail gas produced during the regeneration of the extract-loaded adsorbent contains approximately 11 to 18% of the hydrogen produced (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) and thus represents a loss. This disadvantage does not apply to the process 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.

[0142] Since – apart from electrical consumers such as air compressors – no additional energy is supplied from outside and no excess energy is released to the outside, the H2 product stream must, in the theoretical limiting case, have the same calorific value as the liquid feedstocks methanol or ammonia. Therefore, theoretically, no conversion energy is lost in this process according to the invention. Losses arise only from the fact that the discharged streams are warmer than the input streams, from heat dissipation to the environment via the apparatus walls, and from the mechanical power loss of the air conveying device.

[0143] Assessment parameters:

[0144] The external energy balance in the process according to the invention is determined exclusively by the energy stored in the incoming and outgoing streams (Figure 10). For the theoretical ideal case, in which the incoming streams of methanol / water or ammonia, as well as air and electricity for air compression, import the same amount of power into the process as the outgoing streams of H2 product and exhaust gas export, this process results in an energy efficiency of 100%.

[0145] The energy efficiency can be determined by two approaches, whereby only the second approach can provide an instruction for achieving a low energy efficiency: 1. Approach:

[0146] In the first approach, the chemical energy content of the feed stream and the H2 product stream are used to determine the energy efficiency. The chemical energy content results from the respective lower heat values ​​(LHV) multiplied by the respective mass flows.

[0147] The chemical energy content of the feed stream EF.LHV is calculated as follows from the mass flow mp and the calorific value LHVF

[0148] EF.LHV = mp * LHVF

[0149] The energy content of the H2 product stream EH2 is calculated from the mass flow rriH2 and the calorific value LH H2

[0150] EH2,LHV = mH2 * LH H2

[0151] The chemical conversion losses AEv result from the difference between EF.LHV and EH2,LHV

[0152] AEv, brazen = EF.LHV — EH2,LHV

[0153] For this approach, knowledge of the calorific values ​​LHV (Lower Head Value) for the feed (MeOHfi ÜS si g and NH3fi ü ssi g ) and the product (H2 gas ) are required. These are known from the literature and can be determined as shown in Figure 11. In addition, the mass flow for the energy source used, mp, and the mass flow for the energy source obtained, H2, rriH2, must be known.

[0154] The chemical energy efficiency r|EN,chem of the conversion of MeOH or NH3 to H2 can then be defined as follows:

[0155] T]EN,chem= EH2,LHV / EF.LHV = (EF.LHV ' AEv.chem) / EF.LHV = 1 - AEv.chem / EF.LHV

[0156] For the energy conversion losses AEv, ges the gross electrical power P1 gross, which is introduced into the process, must also be taken into account

[0157] AEv, g it — AE V ,chem + P1 gross

[0158] The energy efficiency r|EN, ge The conversion of MeOH or NH3 to H2 can then be defined as follows:

[0159] T]EN, g es= EH2,LHV / (EF.LHV + SPgross) = (EF.LHV ' AEv.tot) / EF.LHV = 1 ' AEv.tot / EF.LHV

[0160] 2. Perspective:

[0161] In the second approach, the energy flows entering and leaving the process across the balance boundary are balanced to determine the energy efficiency (Figure 10).

[0162] The energy flows in the example calculations are calculated as follows: Gaseous flows:

[0163] Liquid streams. Efiüssi g — rrifiüssig cp ga s, mittei (TT°) ■ AHv with cp gas, mean = average specific heat of the stream for the temperature range between T° and T. T° is the reference temperature and is set to 25°C in the following calculations. The ambient temperature at which the streams enter the balance area of ​​the process and into which the streams exit is also set to 25°C. This has the advantage that no incoming gas streams need to be taken into account, since for these T - T° = 0, but only the outgoing gas streams, which have a higher temperature T than the ambient temperature T°. For the incoming liquid water streams, only the evaporation enthalpy of the water AHV,H2O needs to be taken into account, since for these too T - T° = 0. The evaporation of MeOH and NH3 is taken into account by using the calorific value of the liquid energy carriers LH VMOOH, nutty and LH VNHS, nutty (see Figure 11) for these streams. The enthalpies of vaporization AHV.MOOH and AHV.NHS are therefore already included in the respective calorific values ​​LHVnussig. The respective energy-bearing feed streams can then be treated like guest streams and thus have the value 0.

[0164] According to the second approach, the conversion losses AEv are calculated as follows (see Figure 10):

[0165] AEv= outgoing energy flows - incoming energy flows

[0166] AEv= EH2 + EAG + SQv - mH2O * (- AHV,H2O) - P1 net

[0167] EH2 = mH2 * CPH2, average * (TH2 — T°)

[0168] EAG = mAG * CPAG, medium * (DAY — T°)

[0169] In order to keep the AEv small, in addition to good heat integration, which is expressed in a small EH2 and a small EAG, the flow pressure loss of the fuel gas should be low, the air conveying device should have a high efficiency and the devices should have good insulation.

[0170] For the electrical energy for the compressor P1, only the net power P ne tto must be taken into account, since only the net power is reflected as the energy component in the product stream. The power loss APv = P1 gross - P1 net is not released into the environment in a usable form.

[0171] For the overall efficiency of the process r|EN, ges However, with regard to electrical energy, the gross power P1 must be taken into account. The total energy conversion losses AEv, ges are then

[0172] AEv, g it — AEv + P1 gross

[0173] This results in the overall efficiency of the process r|EN, ge s

[0174] T]EN, g es = 1 - AEv, g es / EF.LHV and the specific total conversion losses AEv.spez related to the H2 product mass flow rriH2

[0175] AEv.spez— AEv.ges / mp|2

[0176] It can be seen that in terms of small energy conversion losses AEv, ges Not only the temperature difference between the outgoing streams and the environment should be kept small, but also the corresponding mass flows. With a constant H2 product flow, this is only the exhaust gas flow. Furthermore, the incoming water flows should be kept small, and as many heat flows as possible should be utilized in the process, and the equipment should be well insulated so that the heat loss flows ZQv are also small. Furthermore, a high mechanical efficiency in the flow machines reduces the energy conversion losses AEv. ges .

[0177] Based on this finding, the task is to develop a process for the conversion of MeOH and NH3 to hydrogen that has the smallest possible enthalpies for the outgoing streams and does not generate any heat loss streams.

[0178] In concrete terms, this means connecting the flows with regard to their heat exchange in such a way that the difference between the two sums of incoming and outgoing specific energy flows, based on the calorific value of the hydrogen product flow, is between 1 and 5 kWh / kg hydrogen.

[0179] The process according to the invention fulfills this task. The process according to the invention enables the overall energy efficiency r|EN, ges of advantageously 80 to 99%, preferably 90 to 98%. The process according to the invention has total energy conversion losses AEv, gesfor the methanol case from 2 to 8 kWh / kg H2, preferably 2.5 to 4 kg / kg H2 and in the ammonia case from 0.5 to 3 kWh / kg H2, preferably 1.0 to 2 kg / kg H2.

[0180] Table 5 lists the terms used again in a table:

[0181] Table 5: Correspondence of the names used in the text for energy flows with the designations used in the figures.

[0182] 1. Example - Methanol:

[0183] Figure 12 shows an example of the process according to the invention for the output of 1000 kg H2 / h according to variant 2 as determined on the basis of a model calculation.

[0184] The example is the result of a thermodynamic simulation with regard to quantities and energies using a BASF internal simulator similar to the simulation program Aspen Plus.

[0185] The example is calculated without heat losses through the process apparatus walls.

[0186] According to the process according to the invention, 6505 kg of methanol and 3723 kg of water at a temperature of 25°C and a pressure of 3 bar must be fed into the process per hour.

[0187] By flowing the liquid feedstocks and the warm reformate in countercurrent through the second reformate cooler, the liquid feedstock mixture is preheated and the reformate is cooled to the adsorption temperature of 40 °C. This requires a heat transfer capacity of 912 kW. The liquid feedstock mixture has a boiling point of 116 °C at 3 bar. 10,227 kg of raw condensate steam are fed to the reformer as reformer feed, and 1 kg / h is fed to the burner as a control stream.

[0188] In the reformer, the raw condensate steam is heated to a reaction temperature of 250°C and catalytically reformed to 70.8 vol% H2, 3.8 vol% CO, and 21.0 vol% CO2. The MeOH equilibrium conversion at 250°C and 3 bar is theoretically 99.9%. The reformate also contains 4.3 vol% unreacted H2O and 0.1 vol% unreacted MeOH. 3750 kW of thermal energy is required for the reforming.

[0189] The reformate is then cooled to 167°C in the first reformate cooler. In turn, the raw condensate vapor is heated from 116°C to 230°C. This heating requires 608 kW of thermal power.

[0190] In the adsorber, 9227 kg / h of extract are adsorbed at 3 bar and 40°C, or stored in the gaseous interstitial volume of the adsorber particles. As stated in Syed Naqvi, "Hydrogen Production," PEP Report 32C, SRI Consulting, September 2007, the example is calculated with 14% H2 losses. This means that 163 kg / h of hydrogen are lost through purge, and 1000 kg / h of hydrogen leaves the adsorber as a 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.

[0191] The extract stream resulting 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.

[0192] In the desorption step, the extract stream is absorbed by a heated air stream and then, after compression, burned in the burner.

[0193] This requires 100,564 kg / h of air, which is first heated to 28 °C in the product cooler or air heater and then to 134 °C in the exhaust gas cooler before entering the desorber to collect the extract. The extract-laden air leaves the desorber as tail gas at a rate of 109,791 kg / h and has a composition of 2.1 vol% H2, 19.2 vol% O2, 72.5 vol% N2, 0.8 vol% CO, 4.5 vol% CO2, 0.9 vol% H2O, and 0.01 vol% MeOH. The desorption step requires heat. Therefore, the temperature in the desorber drops to 122 °C.

[0194] To overcome all flow losses, the tail gas is compressed to 1.2 bar in an air conveyor. The calculations are based on a mechanical efficiency of 90% and an isentropic efficiency of 75%. Compression requires 962 kW of electrical power. During compression, the temperature of the tail gas increases from 122°C to 149°C.

[0195] The compressed tail gas is combusted in a burner together with 1 kg / h of control flow from the evaporator. This produces a hot fuel gas with a temperature of 395°C. This fuel gas is fed into the reformer, where it provides a thermal output of 3750 kW for the reforming reaction and the heating of the gaseous reformer feed from 230 to 250°C.

[0196] The warm fuel gas from the reformer has a temperature of 281°C and is cooled to 154°C in the evaporator before being further cooled to the exhaust gas temperature of 58°C in counterflow to the supplied ambient air.

[0197] The energetic efficiency of the overall process r|En, ges is therefore

[0198] T]EN,ges — 1 - AEv.ges / — 1 (AEv, chem + EAPv) = 1 - [(36053 - 33320) + 962] / 36053 = 0.898.

[0199] The chemical conversion losses AEv, chem are given by: , .

[0200] The specific total conversion losses AEv, spez related to the H2 product mass flow rriH2 are then

[0201] AEv.spez — AEv.ges / m H2 = 3695 kWh / kg H2

[0202] The overall energy efficiency, without taking into account heat losses through the apparatus walls, is r|EN,ges= 90.0%.

[0203] 2. Example - Ammonia

[0204] The example according to variant 2 is shown in Figure 13:

[0205] The example is the result of a thermodynamic simulation with regard to quantities and energies using a BASF internal simulator similar to the simulation program Aspen Plus.

[0206] The example is calculated without heat losses through the process walls. According to the process according to the invention, 7110 kg of liquid ammonia at a temperature of 25°C and a pressure of 10 bar must be fed into the process per hour.

[0207] By flowing the liquid ammonia and the warm reformate in countercurrent in the second reformate cooler, the liquid ammonia is preheated to 92 °C and the reformate is cooled to the adsorption temperature of 40 °C. This requires a heat transfer capacity of 680 kW. The liquid ammonia has a boiling point of 25 °C at 10 bar. 7109 kg of ammonia vapor are fed to the reformer as reformer feed, and 1 kg / h is fed to the burner as a control stream.

[0208] In the reformer, the ammonia vapor is heated to a reaction temperature of 400°C and catalytically reformed to 71.9 vol% H2 and 24.0 vol% N2. The theoretical NH3 equilibrium conversion at 400°C and 10 bar is 92%. The reformate also contains 4.1 vol% unreacted NH3. 5681 kW of thermal energy is required for the reforming.

[0209] The reformate is then cooled to 144°C in the first reformate cooler. In turn, the raw condensate vapor is heated from 25°C to 380°C. This heating requires 1702 kW of thermal power.

[0210] At 10 bar and 40°C, 6109 kg / h of extract are adsorbed in the adsorber, or stored in the gaseous interstitial volume of the adsorber particles. As stated in Syed Naqvi, "Hydrogen Production," PEP Report 32C, SRI Consulting, September 2007, the example assumes 14% H2 losses. This means that 163 kg / h of hydrogen are lost through purge, and 1000 kg / h of hydrogen leaves the adsorber as a 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.

[0211] The extract stream resulting from the adsorber contains 26.4 vol% H2, 62.9 vol% N2 and 10.7 vol% NH3.

[0212] In the desorption step, the extract stream is absorbed by a heated air stream and then, after compression, burned in the burner.

[0213] This requires 21,715 kg / h of air, which is first heated to 56 °C in the product cooler or air heater and then heated to 131 °C in the exhaust gas cooler before entering the desorber to collect the extract. The extract-laden air leaves the desorber as tail gas at a rate of 27,824 kg / h and with a composition of 7.6 vol% H2, 14.9 vol% O2, 74.4 vol% N2, and 3.1 vol% NH3.

[0214] The desorption step requires heat, so the temperature in the desorber drops to 82°C.

[0215] To overcome all flow losses, the tail gas is compressed to 1.2 bar in an air conveyor. The calculations are based on a mechanical efficiency of 90% and an isentropic efficiency of 75%. Compression requires 241 kW of electrical power. During compression, the temperature of the tail gas increases from 82°C to 107°C.

[0216] The compressed tail gas is combusted in a burner together with 1 kg / h of control flow from the evaporator. This produces a hot fuel gas with a temperature of 1000°C. This fuel gas is fed into the reformer, where it delivers a thermal output of 5681 kW for the reforming reaction and the heating of the gaseous reformer feed from 380 to 400°C. The warm fuel gas from the reformer has a temperature of 410°C and is cooled to 274°C in the evaporator before being further cooled to the exhaust gas temperature of 222°C in countercurrent to the supplied ambient air.

[0217] The energetic efficiency of the overall process r|En, ges is therefore

[0218] T]EN,tot = 1 - AEv.tot / EF = 1 — (AEv.chem + LAPv) / EF = 1 — [(34853 — 33320) + 241 ] / 34853 = 0.949.

[0219] The chemical conversion losses AEv.chem are given by:

[0220] AEv.chem = Ev,H2 + Ev.AG + LQv - (-mH2O * AHV,H2O) - SPnetto

[0221] = 60 + 1684 + 0 + 0 - 0.9 * 241 = 1527 kW

[0222] AEv.ges — AEv.chem + SPgross = 1527 + 241 = 1768 kW

[0223] From this, the energetic efficiency of the entire process is calculated

[0224] T|EN,ges = 1 - AEv.ges / EF = 1 — 1768 / 34853 = 0.949

[0225] The specific total conversion losses AEv.spez related to the H2 product mass flow rriH2 are then

[0226] AEv.spez — AEv.ges / m H2 = 1768 kWh / kg H2

[0227] The total energy efficiency without taking into account the heat losses through the apparatus walls is r|EN,ges = 94.9%.

[0228] Variants:

[0229] There are several options for heat integration. What they all have in common is that the fuel gas from the burner is first passed through the heat exchanger in the reformer (A4) and then through the heat exchanger in the evaporator (A2).

[0230] In variant 1 (Figure 4), the vaporized feedstocks are heated in the first reformate cooler before entering the reformer, and the warm reformate is cooled in the second reformate cooler by heating the preheated air. The remaining cooling of the fuel gas is achieved by preheating the feed stream.

[0231] In variant 2 (Figure 5), as in variant 1, the vaporized feedstocks are heated in the first reformate cooler before entering the reformer. However, the warm reformate is cooled not by the heated air, but by the incoming liquid feedstocks. The remaining cooling of the cooled fuel gas after the vaporizer is achieved by heating the heated air.

[0232] In variant 3 (Figure 6), the evaporator serves as the first reformate cooler; the air is heated in the second reformate cooler before entering the desorber, and the warm fuel gas is cooled to the exhaust gas temperature in the feed preheater before leaving the process. In variant 4 (Figure 7), as in variant 3, the evaporator serves as the first reformate cooler; the air is heated in the second reformate cooler before entering the desorber. The warm reformate is cooled by the incoming liquid feedstocks, as in variant 2. The residual cooling of the cooled fuel gas after the evaporator is achieved by heating the heated air.

[0233] Table 7 compares the results for the individual variants.

[0234] The designations are as follows: MeOH-2, for example, means variant 2 for the methanol case. Variant NH3-1 means variant 1 for the ammonia case. NH3-2-00 means variant 2 for ammonia as an energy carrier in the case where the temperature difference between the hot and cold streams in the heat exchangers is zero. However, this would require infinitely large heat exchanger surfaces. Therefore, this variant, with an overall efficiency r|en,ges of 98.1%, represents the theoretical limiting case.

[0235] With higher H2 losses via the tail gas, the gas mixture in the desorber may be explosive. In this case, recirculating a portion of the exhaust gas to the desorber using a cycle gas compressor is recommended (variant NH3-2-KG, Figure 14).

[0236] The influence of H2 losses via the tail gas on the overall efficiency r|en was also investigated. ges. The base case assumes a 14% H2 loss. The cases with 5%, 8%, and 11% were also calculated. While reducing H2 loss from 14% to 11% still results in a significant increase in overall efficiency from 94.9% to 97.1%, a further reduction barely improves overall efficiency.

[0237] An important result of these model calculations is therefore the instruction to reduce the H2 loss via the tail gas to about 11%, which is possible with the process according to the invention, since here the desorption is not carried out by applying a negative pressure as described in the prior art, but an unloaded gas stream (circulating gas or combustion air) can be used to reduce the partial pressure of the extract substances in the gas phase (tail gas).

[0238] Combining the process according to the invention with a fuel cell (FC) for electricity generation offers advantages. This will be explained using the ammonia case (see variant NH3-2-FC, Figure 15):

[0239] • If the fuel cell is operated with a mixture of H2 and N2 rather than pure H2, the driving potential for power generation is reduced at the same pressure, but this can be compensated for by a 25% pressure increase. The associated lower NH3 conversion in the reformer is not significant in this process, as the unconverted NH3 is converted into heat in the burner, which is necessary for the operation of the reforming and evaporation processes.

[0240] • The adsorption effort is significantly reduced, as the separation of NH3 from H2-containing gases requires significantly less effort than the separation of N2 and H2. This is because the boiling point difference between NH3 and H2 is significantly greater than that between N2 and H2. Furthermore, the alkaline effect of NH3 can be utilized very effectively.

[0241] • The complex separation between H2 and N2 occurs naturally in the BSZ through the membrane, which is only permeable to H+ ions.

[0242] • Since the exhaust gases from the fuel cell stack are virtually oxygen-free, they are particularly suitable for desorption. In this case, no recycle gas is required to avoid the explosion zone. • Instead of the usual two air compressors, one for the inventive process and one for the fuel cell stack, only one air compressor is needed. A larger air compressor is always more cost-effective than two small ones. Furthermore, the efficiency increases with the size of the compressor.

[0243] • The use of the exhaust gas streams from the fuel cell as purge gas in the desorption process further improves efficiency. At 95.2%, it is higher than the efficiency of variants with the same heat integration. In comparison, the NH3-2 variant achieves an efficiency of 94.9%, and the NH3-2-KG cycle gas variant achieves an efficiency of 94.8%.

[0244] Table 7: Comparison of the variants for a H2 product flow of 1 kg / h. The enthalpy of the H2 product flow H2 = 33.32 kW

[0245] Comparison of the NH3-2 variant with the state of the art:

[0246] A comparison of the energy efficiency of the process according to the invention with the state of the art provides an overview of the energetic and thus economic advantages of the invention:

[0247] GB 1,079,660 65%

[0248] WO 2018 / 235059 A 1 <78%

[0249] L. Lin et al. <80%

[0250] WO 02 / 071451 A2 85% Lamb et al. 90% EP 3,028,990 >90% (product here is only a gas mixture of H2 / N2 / NH3) Invention 94 to 97% The process according to the invention for producing pure hydrogen has the highest energy utilization rate.

Claims

Patent claims 1. A process for the production of hydrogen from methanol or ammonia, characterized in that methanol or ammonia is evaporated in a first step, reformed in a second step to a hydrogen-containing gas mixture, in a third step the gaseous reformate is cooled to 25 to 100°C, and in a fourth step the hydrogen is separated from the cooled gaseous reformate at a pressure of 1 to 60 bar and a temperature of 25 to 100°C by a sorption process, wherein, in parallel to the first four steps, in a fifth step air is compressed and preheated with the separated hydrogen after the sorption process, in a sixth step the adsorbent loaded with the extract is regenerated with the preheated air, and in a seventh step the extract separated from the adsorbent is combusted with the air, wherein the combustion gases are passed through at least two different heat exchangers,in order to provide, in the flow direction of the fuel gases, (i) first the reaction heat for the reforming of the methanol or ammonia and (ii) subsequently the heat of vaporization for the vaporization of the reformer feed.

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

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

4. Process according to at least one of claims 1 to 3, characterized in that the reformate further heats the preheated air in a heat exchanger before entering the sorption step.

5. A method according to at least one of claims 1 to 3, characterized in that the fuel gases finally heat the preheated air further as step (iii).

6. Process according to at least one of claims 1 to 3, characterized in that the reformate preheats the feed in front of the evaporator in a heat exchanger before entering the sorption step.

7. Process according to claim 6, characterized in that the fuel gases are passed through at least three different heat exchangers in order to (i) first provide the reaction heat for reforming the methanol or ammonia, (ii) then provide the heat of vaporization for vaporizing the reformer feed, and (iii) finally preheat the air for regeneration in the flow direction of the fuel gases.

8. Process according to at least one of claims 1 to 3, characterized in that the fuel gases are passed through at least three different heat exchangers in order to (i) first provide the reaction heat for reforming the methanol or ammonia, (ii) then provide the heat of vaporization for vaporizing the reformer feed, and (iii) finally preheat the feed upstream of the vaporizer in the flow direction of the fuel gases.

9. The process according to claim 8, characterized in that the reformate provides additional heat of vaporization for vaporizing the reformer feed in a heat exchanger before entering the sorption step.

10. Process according to at least one of claims 1 to 3, characterized in that the reformate provides additional heat of vaporization for vaporizing the reformer feed in a heat exchanger before entering the sorption step and preheats the feed before the vaporizer in a further heat exchanger.

11. The process according to claim 10, characterized in that the fuel gases are passed through at least three different heat exchangers in order to (i) first provide the reaction heat for reforming the methanol or ammonia, (ii) then provide the heat of vaporization for vaporizing the reformer feed, and (iii) finally further heat the preheated air in the flow direction of the fuel gases.

12. A process for producing electricity from methanol or ammonia, characterized in that methanol or ammonia is evaporated in a first step, reformed in a second step to a hydrogen-containing gas mixture, in a third step the gaseous reformate is cooled to 25 to 200°C, in a fourth step the hydrogen and, in the case of ammonia, together with the nitrogen, are separated from the cooled gaseous reformate at a pressure of 1 to 60 bar and a temperature of 25 to 200°C by a sorption process, and further electricity is produced from the separated hydrogen in a fuel cell, wherein, in parallel to the first steps, air is compressed and preheated in a fifth step, in a sixth step the adsorbent loaded with the extract is regenerated with the off-gases from the anode and / or cathode side of the fuel cell, and in a seventh step the extract separated from the adsorbent is combusted with the air,wherein the fuel gases are passed through at least three different heat exchangers in order to provide, in the flow direction of the fuel gases, (i) first the reaction heat for the reforming of the methanol or ammonia and (ii) subsequently the heat of vaporization for vaporizing the reformer feed, and (iii) finally to preheat the air and / or the extract separated from the adsorbent.

13. Process according to claim 12, characterized in that ammonia is used as feed.

14. The method according to claim 12 or 13, characterized in that in the sixth step the adsorbent loaded with the extract is regenerated with the off-gases from the anode and cathode sides of the fuel cell.

15. Process according to at least one of claims 12 to 14, characterized in that 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.

Images