METHOD FOR PRODUCING METHANE-CONTAINING FLUID AND APPARATUS FOR PRODUCING METHANE-CONTAINING FLUID

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By employing a series of adiabatic reactors with intermediate cooling for ammonia decomposition and methanation, the method addresses reactor complexity and cost issues, achieving efficient methane production from ammonia and carbon dioxide.

JP2026038393APending Publication Date: 2026-03-06JGC HLDG CORP

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Patent Information

Application Number: JP2024141801
Authority / Receiving Office: JP · JP
Patent Type: Applications
Current Assignee / Owner
Filing Date: 2024-08-23
Publication Date: 2026-03-06

AI Technical Summary

Technical Problem

Existing methods for producing methane from ammonia and carbon dioxide face challenges such as complex reactor structures, temperature imbalances affecting reaction efficiency, and increased costs due to separate reactors and additional equipment, which hinder efficient methane production.

Method used

A method involving a series of adiabatic reactors, where ammonia decomposition and methanation reactions occur in separate stages with intermediate cooling to maintain optimal temperature conditions, using catalysts like Ru and Ni to achieve high conversion rates and efficient methane production.

Benefits of technology

This approach allows for efficient production of methane by maintaining equilibrium conditions across multiple reactors, reducing equipment complexity and costs while achieving high methane yields.

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Abstract

A technology for efficiently producing methane-containing fluids by ammonia methanation is provided. [Solution] In producing a methane-containing fluid from a feed fluid containing ammonia and carbon dioxide, the feed fluid is supplied to a first reactor 31 packed with a first catalyst having ammonia decomposition activity and methanation activity, and ammonia decomposition and methanation proceed to obtain an intermediate product fluid. Next, the intermediate product fluid flowing out of the first reactor 31 is cooled and then supplied to a second reactor 32 packed with a second catalyst having methanation activity, where methanation proceeds to obtain a methane-containing fluid.

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Description

[Technical Field]

[0001] The present invention relates to a technique for producing a methane-containing fluid from a raw material fluid containing ammonia and carbon dioxide. [Background technology]

[0002] Approximately 50-60% of Japan's final energy consumption is heat demand, with natural gas (city gas) accounting for the majority of this. Within this demand, there is a wide range of gas demand, particularly in the high-temperature sector, that is difficult to convert to electricity or other energy sources. To address this issue, the introduction of carbon-neutral (CN) methane, which synthesizes methane (CH4) using carbon dioxide (recovered CO2) captured from combustion exhaust gas and other sources as raw materials, and hydrogen (H2), is being considered. Examples of hydrogen sources used in the synthesis of CN methane include H2 and ammonia (NH3) produced using renewable energy (hereinafter also referred to as "renewable energy").

[0003] Ammonia methanation, which uses ammonia as a hydrogen source to produce methane, involves reacting the hydrogen obtained by decomposing ammonia with carbon dioxide to obtain methane. For example, Patent Document 1 describes a reactor in which an ammonia decomposition chamber containing an ammonia decomposition catalyst and a methanation reaction chamber containing a methanation catalyst are separated by a hydrogen separation membrane. In this reactor, the ammonia decomposition chamber and the methanation reaction chamber are adjacent to each other, so that the heat generated by methanation is utilized to absorb the heat generated by ammonia decomposition. In addition, the equilibrium constraints on ammonia decomposition are alleviated by removing hydrogen from the ammonia decomposition chamber through the hydrogen separation membrane.

[0004] On the other hand, the technology described in Patent Document 1 has the problem of a complicated reactor structure. Furthermore, by placing the space in which the ammonia decomposition reaction proceeds and the space in which the methanation catalytic reaction proceeds adjacent to each other, the reaction conditions on one side are likely to affect the other side. For example, ammonia decomposition is an endothermic reaction, and the higher the reaction temperature, the more likely it is to proceed. On the other hand, methane production, which is an exothermic reaction, is more likely to proceed at lower temperatures. Therefore, temperature conditions that favor ammonia decomposition may not be suitable for efficient methane production.

[0005] As a countermeasure to this problem of a decrease in methane yield as the temperature in the methanation reaction chamber increases, Patent Document 1 describes a method of adjusting the amount of CO2 supplied to prevent the temperature from rising too much, or of resupplying some of the produced methane to the methanation reaction chamber to dilute it. However, methods such as restricting the amount of CO2 supplied or diluting the methane can cause problems such as reduced production efficiency and increased separation and purification costs. On the other hand, resupplying some of the methane requires the installation of piping for the resupply, which increases equipment costs.

[0006] Furthermore, in the case of the reactor structure described in Patent Document 1, the rate at which hydrogen produced in the ammonia decomposition chamber moves to the methanation reaction chamber through the hydrogen separation membrane may become rate-limiting. In this case, some of the hydrogen produced in the ammonia decomposition chamber may be discharged from the reactor without being able to move to the methanation reaction chamber.

[0007] Patent Document 2 describes a methanation reactor in which multiple reactors are connected in series in multiple stages to perform a methanation reaction in which hydrogen is supplied to a feed gas containing carbon dioxide to produce a gas containing methane. These reactors are configured as multi-tubular reactors in which catalysts are packed into multiple reaction tubes. A heat transfer medium, temperature-controlled by a heater or cooler, flows through the body of the reactor housing the multiple reaction tubes and circulates between the reactors, maintaining a constant temperature within each reactor. Furthermore, the gas inlet temperature is adjusted using heaters or coolers installed on the inlet side of each reactor, and the amount of hydrogen supplied to the first and second reactors is adjusted based on the temperature measurement results within the first reactor, thereby achieving a reaction that increases the methane content in the gas composition in line with chemical equilibrium. Patent Document 2 also describes a configuration in which water vapor is introduced along with the feedstock to mitigate temperature increases within the catalyst layer. However, using equipment with a special structure such as a multi-tubular reactor increases the procurement and maintenance costs of the equipment itself. Furthermore, additional equipment such as the heat medium circulation line and steam supply line mentioned above is also required, which also increases the cost of the entire apparatus.

[0008] Next, Patent Document 3 describes a methane production apparatus in which cooling sections are provided between multiple reactors containing a catalyst that promotes a methanation reaction when producing methane using hydrogen and carbon dioxide as feedstocks. These cooling sections cool the product gas generated in the reactor at the upstream stage to a temperature equal to or higher than the temperature at which the methanation reaction starts and lower than the temperature at which the methanation reaction reaches equilibrium and stops. The technology described in Patent Document 3 discloses a method of promoting the methanation reaction by disrupting the equilibrium state by providing a water removal section that removes water from the cooled product gas downstream of one of the cooling sections.

[0009] However, the technologies described in Patent Documents 2 and 3 are technologies for producing methane through a methanation reaction between carbon dioxide and hydrogen, and do not disclose a method for efficiently promoting the decomposition of ammonia, which is an endothermic reaction, and the methanation reaction, which is an exothermic reaction. [Prior art documents] [Patent documents]

[0010] [Patent Document 1] Japanese Patent Application Laid-Open No. 2013-95681 [Patent Document 2] Patent No. 5802551 [Patent Document 3] Japanese Patent Application Laid-Open No. 2015-107942 Summary of the Invention [Problem to be solved by the invention]

[0011] The present invention has been made under such circumstances, and provides a technology for efficiently producing a methane-containing fluid by ammonia methanation. [Means for solving the problem]

[0012] The method is a method for producing a methane-containing fluid from a feed fluid containing ammonia and carbon dioxide, comprising the steps of: supplying the raw material fluid to a first reactor packed with a first catalyst having ammonia decomposition activity and methanation activity, and proceeding with ammonia decomposition and methanation to obtain an intermediate product fluid; cooling the intermediate product fluid exiting the first reactor; and supplying the cooled intermediate product fluid to a second reactor packed with a second catalyst having methanation activity, and proceeding with methanation to obtain the methane-containing fluid.

[0013] The method may include: (a) The first reactor and the second reactor are configured as adiabatic reactors. (b) in the step of cooling the intermediate product fluid, the entire amount of the intermediate product fluid flowing out from the first reactor is cooled, and in the step of obtaining the methane-containing fluid in the second reactor, only the entire amount of the intermediate product fluid cooled in the step of cooling the intermediate product fluid is supplied to the second reactor. (c) The temperature of the raw material fluid supplied to the first reactor is within a range of 300 to 600°C. (d) mixing ammonia and carbon dioxide at a temperature of 100°C or higher to obtain the raw material fluid; and supplying the raw material fluid to the first reactor while maintaining the temperature of 100°C or higher. (e) The difference in temperature of the intermediate product fluid at the outlet of the first reactor relative to the temperature of the raw material fluid at the inlet of the first reactor is within a range of −50 to 150° C., and the conversion rate of ammonia contained in the raw material fluid is 95% or more. (f) The temperature of the intermediate product fluid fed to the second reactor is within the range of 150 to 500°C.

[0014] (g) cooling the methane-containing fluid flowing out of the second reactor as a second intermediate product fluid, and supplying the cooled second intermediate product fluid to a third reactor packed with the second catalyst and proceeding with methanation to obtain the methane-containing fluid, wherein the third reactor is configured as an adiabatic reactor. (h) In (g), in the step of cooling the second intermediate product fluid, the entire amount of the second intermediate product fluid flowing out from the second reactor is cooled, and in the step of obtaining the methane-containing fluid in the third reactor, only the entire amount of the second intermediate product fluid cooled in the step of cooling the second intermediate product fluid is supplied to the third reactor. (i) In (g), the temperature of the second intermediate product fluid supplied to the third reactor is within the range of 150 to 450°C. [Effects of the Invention]

[0015] According to the present invention, the intermediate product fluid obtained by advancing the decomposition of ammonia and methanation in the first reactor is cooled and then supplied to the second reactor, where methanation proceeds, thereby enabling efficient production of a methane-containing fluid. [Brief explanation of the drawings]

[0016] [Figure 1] This is an example of a scheme for producing CN methane from carbon dioxide using renewable energy. [Figure 2] 1 is a graph showing the relationship between temperature and equilibrium composition in ammonia methanation. [Figure 3] 1 is a graph showing the relationship between reaction temperature and raw material conversion rate in ammonia methanation. [Figure 4] 1 is a graph showing the relationship between temperature and equilibrium conversion of ammonia in ammonia methanation. [Figure 5] 1 is a graph showing the relationship between reaction temperature and methane yield in ammonia methanation. [Figure 6] 1 is a configuration example of a methane production apparatus according to an embodiment. [Figure 7] 1 is a first graph showing experimental results of ammonia methanation. [Figure 8] 2 is a second graph showing experimental results of ammonia methanation. DETAILED DESCRIPTION OF THE INVENTION

[0017] (The necessity of ammonia methanation) First, we will explain the need for carbon-neutral (CN) methane production through ammonia methanation, with reference to Figure 1. As shown in Figure 1(a), the conventional method for producing CN methane has been to produce hydrogen (H2) derived from renewable energy in overseas regions suitable for the use of renewable energy, and synthesize it through a methanation reaction (Sabatier reaction) with captured carbon dioxide (CO2). The synthesized CN methane is liquefied, transported by tanker, and supplied to consumers via pipelines at LNG terminals.

[0018] However, the above scheme faces technical challenges, such as reducing the cost of synthesizing CN methane and scaling up the process. Furthermore, because the captured CO2 used in this method is used to synthesize CN methane overseas, CO2 emissions must be managed across borders. Furthermore, producing CN methane overseas requires maintaining a system that accommodates long transportation routes, such as deploying tankers and operating LNG terminals, making it difficult for small and medium-sized businesses to implement.

[0019] Meanwhile, efforts are underway to establish a system for importing fuel ammonia as a relatively inexpensive clean energy source in the near future. As shown in Figure 1(b), fuel ammonia is produced from renewable energy-derived H2, for example by the Haber-Bosch process, and after liquefaction, it is imported as a power generation fuel for ammonia combustion or for co-combustion with other fuels.

[0020] Therefore, if some of this ammonia is purchased and used as a hydrogen source for CO2 methanation, it would be much easier to supply CN methane than transporting CN methane produced overseas. Furthermore, this scheme synthesizes CN methane domestically, allowing for the use of domestically captured CO2 and facilitating CO2 emission management. Furthermore, in this case, ammonia decomposition and CO2 methanation can be carried out in a single facility. As explained below, the ammonia decomposition reaction is highly endothermic, while CO2 methanation is highly exothermic. Therefore, if the facilities for these processes are installed in different countries, each reactor would have a complex configuration. To address this issue, performing these reactions in a single reactor (the first reactor 31 described below) has the advantage of producing only mild heat and simplifying the reactor structure.

[0021] (Ammonia methanation) Ammonia methanation involves the decomposition reaction of NH3 shown in the following formula (1) and the methanation reaction of CO2 shown in the following formula (2). 8NH3→4N2+12H2ΔH=+360kJ …(1) 3CO2+12H2→3CH4+6H2O ΔH=-495kJ …(2) As shown in each equation, the NH3 decomposition reaction is an endothermic reaction, while the methanation reaction is an exothermic reaction. The balance of these reaction heats changes depending on the thermodynamic equilibrium conditions and the progress of the reaction, which change depending on the reaction pressure and reaction initiation temperature. Therefore, when ammonia methanation is carried out, the temperature in the reaction system may either rise or fall.

[0022] In addition, a catalyst is used in both the reactions of formulas (1) and (2). The catalysts used in these reactions are solid catalysts in which an active metal such as ruthenium (Ru) or nickel (Ni) is supported on a support of a metal oxide such as ceria (CeO2), alumina (Al2O3), or zirconia (ZrO2), or a composite metal oxide containing multiple metal elements. Ru and Ni have catalytic activity for both the reactions of formulas (1) and (2). Furthermore, Ru has a higher reaction activity than Ni.

[0023] Let's take a closer look at the reaction characteristics of ammonia methanation described above. Figure 2 is a graph showing the change in the thermodynamic equilibrium composition of CH4, CO2, and carbon monoxide (CO) versus temperature. From the perspective of thermodynamic equilibrium, Figure 2 shows that the lower the temperature and the higher the pressure, the higher the CH4 yield tends to be. The minimum temperature required for the methanation reaction of formula (2) to proceed is approximately 200°C, for example, 150°C.

[0024] On the other hand, at high temperatures, the ratios of CO and CO2 increase due to constraints of thermodynamic equilibrium, and the yield of CH4 decreases. CO is thought to be produced mainly through the reverse shift reaction (CO2 + H2 → CO + H2O). From this perspective, it is preferable to carry out ammonia methanation at a reaction temperature of, for example, 500°C or less. Furthermore, as mentioned above, the higher the pressure, the higher the yield of CH4 can be, but from the viewpoint of industrial operation of the equipment, it is practical to set the reaction pressure within the range of about 0.1 to 4 MPaG.

[0025] Considering the reaction temperature characteristics of the equilibrium composition shown in Figure 2, it would seem that ammonia methanation would produce better results if it were carried out under lower reaction temperature conditions. However, the actual reaction is affected by catalyst performance. Figure 3 is a graph showing the relationship between reaction temperature and the conversion rates of NH3 and CO2. In the figure, the dashed line is the curve showing the equilibrium conversion rate of NH3, and the solid line is the curve showing the equilibrium conversion rate of CO2. In Figure 3, the equilibrium conversion rate curve for NH3 is stuck near 100%. Therefore, Figure 4 shows an enlarged graph of the relationship between temperature and equilibrium NH3 conversion rate in the region near 100%.

[0026] According to these equilibrium conversion curves for NH3 and CO2, the NH3 conversion rate is close to 100% at reaction temperatures between 300 and 500°C, and the methanation reaction of formula (2) can proceed under conditions where a sufficient amount of H2 is supplied. In this case, from the perspective of suppressing the constraints of thermodynamic equilibrium, performing ammonia methanation at a lower reaction temperature will result in a higher CO2 conversion rate.

[0027] On the other hand, the plot in Figure 3 shows an example of the conversion rate of NH3 and CO2 in a preliminary experiment of ammonia methanation. The preliminary experiment used a catalyst with 3 wt% Ru supported on a CeO2 carrier, and the reaction pressure was 0.7 MPaG, and the WHSV (Weight Hourly Space Velocity) was 2.4 h -1 The reaction temperature (= temperature at which the raw material fluid was supplied to the catalyst layer) was set to four conditions: 400, 450, 500, and 600°C. NH3 and CO2 were supplied at flow rates that achieved the stoichiometric molar ratio based on equations (1) and (2).

[0028] The open circle plots in Figure 3 show the NH3 conversion at each reaction temperature in the preliminary experiment, while the filled circle plots show the CO2 conversion. Looking at these plots, it can be seen that at a reaction temperature of 400°C, where a higher CO2 conversion rate should be obtained according to the equilibrium conversion curve, the CO2 conversion rate remains at around 65%. Furthermore, as shown in Figures 3 and 4, even for NH3, where an equilibrium conversion rate of over 99.8% should be obtained over a wide reaction temperature range from 200 to 700°C, the results of the preliminary experiment at a reaction temperature of 400°C showed a conversion rate of less than 70%.

[0029] The results of these preliminary experiments indicate that the catalytic activity (reaction rate) of the NH3 decomposition reaction at the reaction temperature is a constraint, and that in some cases the equilibrium conversion rate cannot be reached at low reaction temperatures. This is thought to result in a shortage of H2 supplied by the NH3 decomposition reaction, restricting the progress of the methanation reaction and resulting in a low conversion rate for CO2. In particular, activity drops sharply as the reaction temperature is lowered from the aforementioned 400°C, and it is thought that at 300°C it becomes difficult to proceed with the methanation reaction. On the other hand, at the three reaction temperatures (450, 500, and 600°C) where an NH conversion rate close to the equilibrium conversion rate of 95% or more was obtained, it was confirmed that as the reaction temperature increased, the CO conversion rate decreased in line with the equilibrium conversion rate.

[0030] To summarize the points confirmed above using Figures 2 to 4, whether the NH decomposition reaction of Equation (1) reaches equilibrium conversion depends on catalytic activity. Therefore, it is necessary to select a catalyst and reaction temperature that achieve a conversion rate close to equilibrium conversion under the conditions (reaction pressure, WHSV, etc.) in which the catalyst is used. Furthermore, the methanation reaction of Equation (2) proceeds even at low reaction temperatures below 200°C. In particular, from the perspective of obtaining a high methane yield, it is preferable to proceed with the methanation reaction at as low a reaction temperature as possible. However, if the NH decomposition reaction in the preceding stage proceeds and sufficient H2 is not supplied, the progress of the methanation reaction will be suppressed (plot for a reaction temperature of 400°C in Figure 3).

[0031] According to these findings, ammonia methanation should be carried out at a reaction temperature at which the NH decomposition reaction can reach equilibrium conversion, and the methanation reaction should be carried out at a low reaction temperature at which the CO conversion rate is higher. This finding, which is particularly difficult to derive from thermodynamic considerations alone, is that at low reaction temperatures, catalytic activity becomes a constraint and equilibrium may not be reached.

[0032] (Reaction procedure) Based on the kinetic considerations discussed above, we will now consider a reaction procedure suitable for carrying out ammonia methanation. Figure 5 shows the evolution of reaction temperature and CH4 yield calculated from the thermodynamic equilibrium of equations (1) and (2) assuming a reaction in a fixed-bed adiabatic reactor at a reaction pressure of 0.7 PaG. When a mixed gas of NH3 and CO2 is supplied as the feed fluid, the equipment configuration can be simplified if the mixed gas can be supplied to a single reactor packed with a catalyst that has both NH3 decomposition activity and methanation activity, and the desired yield of CH4 can be obtained.

[0033] However, as mentioned above, the temperature in the reaction system (reactor) rises throughout the ammonia methanation process, and the equilibrium yield of CH4 is governed by the reactor outlet temperature. For example, the dashed-dotted line connecting the plots indicated by open triangles in Figure 5 shows the reaction operating line when the temperature of the feed fluid at the reactor inlet (inlet temperature) is set to 300°C. In the case of a single-stage reactor, the temperature rises in the catalyst layer as the reaction progresses, and the temperature at the outlet (outlet temperature) rises to approximately 430°C. In this case, the equilibrium yield of CH4 does not reach 90%. The inventors of this application aim to achieve a CH4 yield (approximately equal to the CO2 conversion rate in the temperature range where CO by-production is low at thermodynamic equilibrium) of 90% or more.

[0034] If we wanted to improve the CH4 yield solely from the viewpoint of thermodynamic equilibrium, we could simply lower the reactor inlet temperature. For example, the reaction operating line (dashed line) connecting the open square plots in Figure 5 shows an example in which the inlet temperature of the feed fluid was set to 200°C. In this example, the outlet temperature was kept below 400°C, resulting in a CH4 yield of nearly 95%.

[0035] However, the ammonia decomposition reaction does not proceed at an inlet temperature of around 200°C. As a result, a sufficient amount of H2 is not supplied, making it difficult to promote the methanation reaction.

[0036] Therefore, in this embodiment, ammonia methanation is carried out using a first reactor in which the NH decomposition reaction and the methanation reaction proceed, and a second reactor and a third reactor in which the methanation reaction proceeds. Then, between the first and second reactors, and between the second and third reactors, fluids flowing out from the upstream reactors ("intermediate product fluid" described below), are cooled, respectively, to achieve a high CH yield across the multiple reactors (first reactor to third reactor).

[0037] The solid line connecting the plots of black circles in Figure 5 corresponds to the reaction operating line carried out in the first to third reactors described above. The important points in this reaction operation are (i) in the first reactor, the reaction is allowed to proceed until the equilibrium composition of the ammonia methanation reaction at the outlet temperature is reached, and (ii) then, in the second and third reactors, the methanation reaction is allowed to proceed at a temperature lower than the outlet temperature of the previous reactor, and the fluids to be treated (intermediate product fluid, second intermediate product fluid) are cooled at the inlet side of each reactor to obtain the target CH4 yield.

[0038] In particular, the technical viewpoint of (i) is difficult to derive from the configuration of a reactor in which the ammonia decomposition chamber and the methanation reaction chamber are separated by a hydrogen separation membrane and NH3 and CO2 are supplied separately, as in Patent Document 1, cited as an example of the background art. Furthermore, the technologies described in Patent Documents 2 and 3, also cited as examples of the background art, are documents that describe only the methanation reaction, not the NH3 decomposition reaction. In particular, as mentioned above, the ammonia decomposition reaction and the CO2 methanation reaction start in different temperature ranges. Therefore, Patent Documents 2 and 3 only describe a multi-stage reactor for the CO2 methanation reaction, and it is not possible to obtain a configuration that optimizes the two different reactions (ammonia decomposition reaction and CO2 methanation reaction). Hereinafter, with reference to FIG. 6, the configuration of an apparatus (ammonia methanation apparatus 1) for carrying out the method for producing a methane-containing fluid according to the embodiment will be described.

[0039] (Ammonia methanation device 1) As shown in Figure 6, the ammonia methanation apparatus (apparatus for producing a methane-containing fluid) 1 of this embodiment includes three reactors: a first reactor 31 in which the decomposition reaction of NH3 and the methanation reaction proceed with respect to the raw material fluid; and a second reactor 32 and a third reactor 33 in which the methanation reaction proceeds with respect to the fluid flowing in from the upstream reactors (first reactor 31, second reactor 32). A first cooler (intermediate product fluid cooling unit) 311 is provided between the first reactor 31 and the second reactor 32 to cool the fluid (intermediate product fluid) flowing out of the first reactor 31. A second cooler (second intermediate product fluid cooling unit) 321 is provided between the second reactor 32 and the third reactor 33 to cool the fluid (second intermediate product fluid) flowing out of the second reactor 32.

[0040] The correspondence between the schematic configuration of the above-described ammonia methanation apparatus 1 and the reaction operation line shown by the solid line in Figure 5 will be described below. In the first reactor 31, a reaction indicated by R1 in the figure proceeds, and in the first cooler 311, a cooling operation indicated by C1 is carried out. Next, in the second reactor 32, a reaction indicated by R2 proceeds, and in the second cooler 321, a cooling operation indicated by C2 is carried out. Finally, in the third reactor 33, a reaction indicated by R3 proceeds.

[0041] The first reactor 31, the second reactor 32, and the third reactor 33 are configured, for example, as fixed-bed reactors. The first reactor 31 is filled with a first catalyst having NH3 decomposition activity and methanation activity. The second reactor 32 and the third reactor 33 are filled with a second catalyst having methanation activity. As described above, both Ru and Ni have both NH3 decomposition activity and methanation activity, but Ru has higher reaction activity than Ni. Therefore, the active metal supported on each catalyst can be selected taking into consideration the type of reaction required in each reactor 31, 32, and 33.

[0042] As explained using FIG. 3, in order to carry out the reaction procedure R1 shown in FIG. 5 in the first reactor 31, it is required to decompose the NH supplied to the first reactor 31 to a conversion rate close to the equilibrium, for example, 95% or more, and further to proceed with the methanation reaction until the equilibrium yield at the outlet temperature of the first reactor 31 is reached. From this viewpoint, the first catalyst may be, for example, a catalyst in which highly active Ru is supported on a carrier. When designing the first catalyst, it is possible to confirm that the catalyst has the activity required to carry out the reaction procedure R1 under the reaction conditions (reaction pressure, WHSV, etc.) in the first reactor 31, for example, by conducting the preliminary experiment explained using FIG. 3.

[0043] On the other hand, the reaction procedures R2 and R3 carried out in the second reactor 32 and the third reactor 33 are required to have at least the activity to proceed with a methanation reaction to the equilibrium yield of CH at the outlet temperature for CO2 (which also contains a trace amount of CO) contained in the fluid cooled in the second cooler 321 and the condenser 331. From this perspective, the second catalyst may be, for example, a catalyst in which Ni, which is less active than Ru but less expensive, is supported on a carrier.

[0044] However, the above-mentioned selection of active metals is merely an example, and it is not essential to use Ru for the first catalyst and Ni for the second catalyst. Even when Ni is used for the first catalyst, it may be possible to carry out reaction operation R1 by increasing the catalyst loading amount in the first reactor 31 to increase the WHSV or by raising the reaction temperature. From this perspective, the use of Ni as the active metal for both the first catalyst and the second catalyst is not excluded. Contrary to this example, it is also possible to use Ru for both the first catalyst and the second catalyst.

[0045] The first reactor 31, the second reactor 32, and the third reactor 33 are configured as, for example, adiabatic reactors. An adiabatic reactor is a reactor in which heat exchange operation using a heat medium or the like supplied from the outside is not performed. Generally, the outer wall surface of an adiabatic reactor is covered with a heat insulating material, but in this example, providing the reactor with a heat insulating material is not an essential requirement. Natural heat dissipation from the reactor to the outside air may occur.

[0046] On the upstream side of the inlet of the first reactor 31, there are provided a CO2 vaporizer 201, which is arranged on the CO2 supply line and vaporizes and heats CO2 supplied, for example, in a liquid state; an NH3 vaporizer 202, which is arranged on the NH3 supply line and vaporizes and heats NH3 supplied, for example, in a liquid state; and a raw material heater 203, which is arranged on the raw material supply line 211 that supplies a raw material fluid, which is a mixed gas of NH3 and CO2, and heats the raw material fluid.

[0047] Mixing NH3 and CO2 may produce ammonium carbamate (CH6N2O2). CO2 vaporizer 201 and NH3 vaporizer 202 heat the respective gases so that NH3 and CO2 are mixed at a temperature higher than the precipitation temperature of ammonium carbamate, for example, 100°C or higher. The mixed raw material fluid is further heated in raw material heater 203 and supplied to first reactor 31 while maintaining the temperature at 100°C or higher.

[0048] The raw material heater 203 heats the raw material fluid to a supply temperature within a range of, for example, 300 to 600°C. The supply temperature of the raw material fluid to the first reactor 31 is selected in consideration of the flow rate of the raw material fluid, the activity of the first catalyst, the amount of the first catalyst packed in the first reactor 31, the reaction pressure, the WHSV, etc., and is a temperature at which the reaction operation R1 described with reference to Fig. 5 can be carried out. For example, in the example of the reaction operation R1 shown in Fig. 5, the supply temperature (the inlet temperature of the first reactor 31) is 450°C. As described above, depending on the heat balance within the first reactor 31, the outlet temperature of the first reactor 31 may be lower than the inlet temperature, as in the reaction procedure R1' also shown in Figure 5. Even in such a case, the ammonia methanation apparatus 1 can be designed based on the same concept as in the example shown below.

[0049] The outlet of the first reactor 31 and the inlet of the second reactor 32 are connected by piping that constitutes a first connection line (intermediate product fluid outflow line) 312. The above-mentioned first cooler 311 is interposed in this first connection line 312. An example of a configuration in which the first connection line 312 is not connected to a recycle line that extracts a portion of the fluid (intermediate product fluid) that has flowed out from the first reactor 31 and returns it to the inlet side of the first reactor 31 is shown. In addition, the first cooler 311 is not configured as a condenser that condenses and removes the water generated in the reaction of formula (2).

[0050] The outlet of the second reactor 32 and the inlet of the third reactor 33 are connected by piping that constitutes a second connection line (second intermediate product fluid supply line) 322. The second cooler 321 described above is interposed in this second connection line 322. An example of a configuration in which the second connection line 322 is not connected to a recycle line that extracts a portion of the fluid (second intermediate product fluid) flowing out of the second reactor 32 and returns it to the inlet side of the second reactor 32 is also not configured as a condenser.

[0051] Furthermore, a pipe constituting a third connection line 332 is connected to the outlet of the third reactor 33. This third connection line 332 is provided with a condenser 331 for cooling the fluid (methane-containing fluid) flowing out of the third reactor 33 and condensing H2O. A gas-liquid separation tank 4 is connected to the third connection line 332 of the condenser 331, and the methane-containing fluid is separated into a gas containing CH4 and N2 and a liquid containing H2O. An example of a configuration in which a recycle line for extracting a portion of the methane-containing fluid and returning it to the inlet side of the third reactor 33 is not connected to the third connection line 332.

[0052] According to the above-described configurations of the first connecting line 312 and the second connecting line 322, the first cooler 311 cools the entire amount of the intermediate product fluid flowing out from the first reactor 31. Then, only the entire amount of the intermediate product fluid cooled by the first cooler 311 is supplied to the second reactor 32. The first cooler 311 cools the intermediate product fluid to a temperature in the range of 150 to 500°C, for example (cooling operation C1). The temperature of the intermediate product fluid after being cooled by the first cooler 311 becomes the supply temperature to the second reactor 32.

[0053] The supply temperature of the intermediate product fluid to the second reactor 32 is selected at a temperature at which the reaction operation R2 described with reference to Fig. 5 can be carried out, taking into consideration the flow rate of the intermediate product fluid, the activity of the second catalyst, the amount of the second catalyst packed in the second reactor 32, the reaction pressure, the WHSV, etc. For example, in the example of reaction operation R2 shown in Fig. 5, the supply temperature (the inlet temperature of the second reactor 32) is 350°C.

[0054] Furthermore, according to the configurations of the second connecting line 322 and the third connecting line 332 described above, the second cooler 321 cools the entire amount of the second intermediate product fluid flowing out from the second reactor 32. Then, only the entire amount of the second intermediate product fluid cooled by the second cooler 321 is supplied to the third reactor 33. The second cooler 321 cools the second intermediate product fluid to a temperature within a range of, for example, 150 to 450°C (cooling operation C2). The temperature of the second intermediate product fluid after being cooled by the second cooler 321 becomes the supply temperature to the third reactor 33.

[0055] The supply temperature of the second intermediate fluid to the third reactor 33 is selected at a temperature at which the reaction operation R3 described with reference to Fig. 5 can be carried out, taking into consideration the flow rate of the second intermediate fluid, the activity of the second catalyst, the amount of the second catalyst packed in the third reactor 33, the reaction pressure, the WHSV, etc. For example, in the example of reaction operation R3 shown in Fig. 5, the supply temperature (the inlet temperature of the third reactor 33) is 250°C.

[0056] (Generation of methane-containing fluids) A method for producing a methane-containing fluid using the ammonia methanation apparatus 1 having the above-described configuration will be described. For example, liquid CO2 and liquid NH3 are separately supplied to the ammonia methanation apparatus 1, vaporized in a CO2 vaporizer 201 and an NH3 vaporizer 202, respectively, and heated to a temperature of 100°C or higher, at which ammonium carbamate does not precipitate. The gaseous CO2 and NH3 are then mixed and supplied to a feedstock heater 203 as a feedstock fluid. The molar ratio of NH3 to CO2 in the feedstock fluid can be controlled to approach the stoichiometric molar ratio (NH3:CO2 = 8:3) derived from the above-described equations (1) and (2). The feedstock fluid is heated in the feedstock heater 203 to a supply temperature within a range of 300 to 600°C, for example, 450°C, and then supplied to the first reactor 31.

[0057] In the first reactor 31, the first catalyst causes the decomposition reaction of NH3 of formula (1) and the methanation reaction of formula (2) to proceed, yielding an intermediate product fluid (reaction operation R1, step of obtaining an intermediate product fluid). In this example, the first reactor 31 decomposes 95% or more of NH3 (NH3 conversion rate of 95% or more). In addition, the methanation reaction is allowed to proceed so that the CO2 conversion rate is, for example, 75% or more. As mentioned above, depending on the overall heat balance of the reactions of formulas (1) and (2), the temperature inside the first reactor 31 may rise or fall. As a result, the temperature difference (ΔT = outlet temperature T out -Inlet temperature T in ) is within the range of -50 to 150°C.

[0058] The entire amount of the intermediate product fluid flowing out of the first reactor 31 is cooled in the first cooler 311 to a temperature in the range of 150 to 500°C, for example, 350°C, which is the supply temperature to the second reactor 32 (cooling operation C1, step of cooling the intermediate product fluid). In the example shown in FIG. 5, conditions are selected such that ΔT in the first reactor 31 is 100°C or less, for example, from the viewpoint of increasing the methane yield while protecting the catalyst. This keeps ΔT in the catalyst layer relatively small, and even when an inexpensive adiabatic reactor is used, it is possible to prevent the catalyst from cracking or pulverizing. The same applies to the second reactor 32 and the third reactor 33 described below. Thereafter, only the entire amount of the intermediate product fluid cooled in the first cooler 311 is supplied to the second reactor 32.

[0059] In the second reactor 32, the methanation reaction of formula (2) proceeds in the presence of the second catalyst, yielding a methane-containing fluid (reaction operation R2, a step of obtaining a methane-containing fluid from the intermediate product fluid). At this time, a decomposition reaction of NH3 remaining in the intermediate product fluid may also proceed. In the second reactor 32 of this example, the methanation reaction is allowed to proceed so that the CO2 conversion rate is, for example, 85% or higher. As described above, the reaction of formula (2) is an exothermic reaction, and therefore the temperature inside the second reactor 32 rises. As a result, the temperature of the methane-containing fluid at the outlet of the second reactor 32 is within a range of 10 to 200°C higher than the supply temperature of the intermediate product fluid at the inlet.

[0060] The entire amount of the methane-containing fluid flowing out from the second reactor 32 is supplied as a second intermediate product fluid to the second cooler 321. In the second cooler 321, the second intermediate product fluid is cooled to a temperature within a range of 150 to 450°C, for example, 250°C, which is the supply temperature to the third reactor 33 (cooling operation C2, a step of cooling the methane-containing fluid as a second intermediate product fluid). Thereafter, only the entire amount of the second intermediate product fluid cooled in the second cooler 321 is supplied to the third reactor 33.

[0061] In the third reactor 33, the methanation reaction of formula (2) proceeds in the presence of the second catalyst, and a methane-containing fluid is obtained (reaction operation R3, a step of obtaining a methane-containing fluid from the second intermediate product fluid). At this time, a decomposition reaction of NH3 remaining in the second intermediate product fluid may also proceed. In the third reactor 33 of this example, the methanation reaction is allowed to proceed so that the CO2 conversion rate is, for example, 90% or more. As described above, the reaction of formula (2) is an exothermic reaction, and therefore the temperature inside the second reactor 32 rises.

[0062] The methane-containing fluid flowing out of the third reactor 33 is cooled in a condenser 331, which condenses the H2O. Thereafter, the methane-containing fluid that has become a gas-liquid mixed fluid is separated into a liquid (H2O) and a gas (CH4, N2) in a gas-liquid separation tank 4. Furthermore, the N2 in the gas is adsorbed and removed, for example, by a PSA (Pressure Swing Absorption) method, and the CH4 from which the N2 has been removed is supplied to consumers as CN methane.

[0063] (effect) According to this embodiment, the decomposition of NH3 and methanation at the outlet of the first reactor 31 are allowed to reach a nearly equilibrium composition under the outlet reaction conditions. Furthermore, the intermediate product fluid and second intermediate product fluid obtained in the first reactor 31 and second reactor 32 are cooled and then supplied to the second reactor 32 and third reactor 33, respectively, to proceed with methanation. During this process, the temperature change in the catalyst layer of each reactor 31, 32, and 33 is kept small, protecting the catalyst and allowing efficient production of a methane-containing fluid. Furthermore, in a configuration using adiabatic reactors 31, 32, and 33, a methane-containing fluid can be produced without using a complicated cooling system.

[0064] (Variation) 1, an example was shown in which three reactors (first reactor 31, second reactor 32, third reactor 33) are provided in the ammonia methanation apparatus 1, and the intermediate product fluid and the second intermediate product fluid are cooled by the first cooler 311 between the first reactor 31 and the second reactor 32 and the second cooler 321 between the second reactor 32 and the third reactor 33, respectively, to the supply temperatures of the downstream reactors (second reactor 32, third reactor 33). However, the number of reactors to be installed is not limited to this example.

[0065] For example, the number of reactors may be changed depending on the processing amount of the raw material fluid, the activity of the catalyst, etc. If at least the first reactor 31 in which the decomposition reaction of NH3 in the raw material fluid and the methanation reaction proceed, the second reactor 32 in which the methanation reaction of the intermediate product fluid proceeds, and the first cooler 311 in which the intermediate product fluid cools, a methane-containing fluid can be produced efficiently.

[0066] Furthermore, the first reactor 31, the second reactor 32, and the third reactor 33 are not limited to being constituted by adiabatic reactors. For example, at least one of the first reactor 31, the second reactor 32, and the third reactor 33 may be constituted by a so-called isothermal reactor. An example of an isothermal reactor is a configuration in which a catalyst layer is formed in one or more reaction tubes, and heat exchange is performed between the catalyst layer and a heat medium flowing outside the reaction tubes, thereby keeping the temperature of the catalyst layer constant.

[0067] In addition, it is not an essential requirement that the entire amount of the intermediate product fluid flowing out of the first reactor 31 be cooled in the first cooler 311 and then supplied to the second reactor 32, and that the entire amount of the second intermediate product fluid flowing out of the second reactor 32 be cooled in the second cooler 321 and then supplied to the third reactor 33. This does not exclude a configuration in which a recycle line is provided, as necessary, to return a portion of the intermediate product fluid cooled in the first cooler 311 to the inlet side of the first reactor 31 or a portion of the second intermediate product fluid cooled in the second cooler 321 to the inlet side of the second reactor 32. As in the above example, it is also not excluded that a recycle line is provided, as necessary, to extract a portion of the methane-containing fluid cooled in the condenser 331 and return it to the inlet side of the third reactor 33. [Example]

[0068] (experiment) Using raw material gas simulating the intermediate product fluid (assumed to have reached equilibrium composition) flowing out from the outlet of the first reactor 31 in the first stage and the outlet of the second reactor 32 in the second stage, methanation reactions were carried out in the second reactor 32 in the second stage and the third reactor 33 in the third stage, and changes in CO2 conversion rate were confirmed. A. Experimental Conditions Test reactors corresponding to the second reactor 32 and the third reactor 33 were constructed using a packed bed of a 3 wt% Ru / CeO2 catalyst or a 50 wt% Ni / Al2O3 catalyst packed into a reaction tube. Using these test reactors, a second-stage methanation reaction corresponding to the second reactor 32 and a third-stage methanation reaction corresponding to the third reactor 33 were carried out under a reaction pressure of 0.7 MPaG. The simulated gas supplied to the second reactor 32 was assumed to have an equilibrium composition after passing through the first reactor 31, and the simulated gas supplied to the third reactor 33 was assumed to have an equilibrium composition after passing through the first reactor 31 and the second reactor 32. In assuming the composition of these simulated gases and conducting experiments on the methanation reactions in the second reactor 32 and the third reactor 33, the WHSV calculated from the sum of the raw material NH3 and CO2 was 2.4 h for the first reactor 31. -1 , 0.2 h per second reactor 32 -1 , 0.1h for the third reactor -1 Under these conditions, in the second-stage methanation reaction, the supply temperature of the intermediate product fluid (first-stage outlet CO2 conversion rate of 83.5%, corresponding to the first reactor 31) was varied between 350, 400, and 430°C. In the third-stage methanation reaction, the supply temperature of the second intermediate product fluid (second-stage outlet CO2 conversion rate of 89.9%, corresponding to the second reactor 32) was varied between 250, 300, and 323°C.

[0069] B. Experimental Results Figure 7 shows the experimental results for the second and third stage methanation reactions using a Ru catalyst, and Figure 8 shows the experimental results for the second and third stage methanation reactions using a Ni catalyst. In each figure, the horizontal axis shows the supply temperature of the intermediate product fluid and the second intermediate product fluid, and the vertical axis shows the CO2 conversion rate based on the feed fluid at the inlet of the first reactor. The equilibrium CO2 conversion rate at each temperature is also shown as a curve. The black circles in the figure show the CO2 conversion rate in the second stage, and the black squares show the CO2 conversion rate in the third stage.

[0070] As can be seen from these figures, the assumed feedstock compositions at the second and third stage inlets reach the equilibrium CO2 composition at each reaction temperature under all conditions tested by passing them through test reactors filled with Ru catalyst and Ni catalyst. [Explanation of symbols]

[0071] 1. Ammonia methanation device 201 CO2 Vaporizer 202 NH3 vaporizer 203 Raw material heater 211 Raw material supply line 31 First Reactor 311 First Cooler 312 First connecting line 32 Second reactor 321 Second Cooler 322 Second connecting line 33 Third Reactor 331 Condenser 332 Third Connection Line 4 Gas-liquid separation tank

Claims

1. A method for producing a methane-containing fluid from a feedstock fluid containing ammonia and carbon dioxide, comprising: supplying the raw material fluid to a first reactor packed with a first catalyst having ammonia decomposition activity and methanation activity, and proceeding with ammonia decomposition and methanation to obtain an intermediate product fluid; cooling the intermediate product fluid exiting the first reactor; and supplying the cooled intermediate product fluid to a second reactor filled with a second catalyst having methanation activity, and proceeding with methanation to obtain the methane-containing fluid. A method for producing a methane-containing fluid.

2. The first reactor and the second reactor are configured as adiabatic reactors.