JP7829396B2 - Method for producing high calorific value fuel gas - Google Patents

Description

Applicable under Article 30, Paragraph 2 of the Patent Law. March 5, 2021. Publication of the proceedings of the 127th Catalysis Symposium on CD and USB.

Applicable under Article 30, Paragraph 2 of the Patent Law; Publicly disclosed through presentation at the 127th Catalysis Symposium on March 17, 2021.

This invention relates to a method for producing a high-calorific value fuel gas mainly composed of methane by methaneating ethanol in the presence of hydrogen.

City gas, primarily composed of natural gas, is an energy source with a low environmental impact compared to other fossil fuels such as petroleum, as it emits less carbon dioxide during combustion. However, in recent years, with environmental problems such as global warming becoming increasingly serious, further reductions in carbon dioxide emissions are desired.

In recent years, biomass fuels such as bioethanol have attracted attention as a new energy resource. Bioethanol is obtained by alcoholic fermentation of sugars, which are produced by saccharifying starch contained in sugarcane juice or grains such as corn. Although carbon dioxide is produced when bioethanol is burned, the plants used as raw materials absorb a corresponding amount of carbon dioxide from the air during their growth process. Therefore, when considering the total amount of carbon dioxide produced during the plant's growth process, it can be said that bioethanol does not increase the overall concentration of carbon dioxide in the atmosphere. For this reason, bioethanol is considered a carbon-neutral fuel.

If a method for efficiently producing fuel gas primarily composed of methane from ethanol can be established, it would be possible to reduce carbon dioxide emissions by, for example, using bioethanol as a raw material for city gas.

Patent Document 1 discloses a method for producing fuel gas by catalytically cracking aliphatic alcohols having 2 to 5 carbon atoms or a mixture of such alcohols and water in the presence of a ruthenium-based catalyst. According to this document, the catalytic cracking reaction is usually carried out under conditions of 400°C to 700°C and 2 to 10 kg/ ^cm² (absolute pressure), and it is shown that by catalytically cracking ethanol in the presence of water vapor, a fuel gas containing 32% to 44% hydrogen, 29% to 40% methane, and 16% to 23% carbon dioxide by volume, with a strength of 4200 to 5050 kcal/ ^m³ , was obtained.

Patent Document 2 discloses a method for gasifying ethanol by mixing it with water in a predetermined ratio and passing it through a nickel-based catalyst at 300°C to 600°C. It also states that by cooling the resulting gaseous mixture to 250°C to 350°C and then introducing it into a catalytic methane process, a fuel gas with a high methane concentration that can be used as synthetic natural gas after carbon dioxide removal can be obtained. However, there is no specific description of the catalytic methane process, nor is there any specific description of the fuel gas composition after the catalytic methane process.

Patent Document 3 discloses a method for producing methane-containing gas from an ethanol-containing raw material containing ethanol and water vapor, the method comprising a reaction step of passing the ethanol-containing raw material once at a reaction temperature of 400°C or higher through a catalyst having rhodium or ruthenium supported on the surface of an inorganic oxide that serves as a carrier.
Even with the method described in this document, the methane concentration in the resulting methane-containing gas is only about 55%, making it unsuitable for use as a raw material for city gas.

Natural gas is commonly used as a raw material for city gas, and its main component is methane, with small amounts of ethane, propane, and butane. Natural gas typically does not contain hydrogen or carbon monoxide, and carbon dioxide is removed during the natural gas refining process. In particular, in the case of city gas produced using liquefied natural gas as a raw material, hydrogen, carbon monoxide, and carbon dioxide are almost completely removed during the liquefaction and refining process, so they are practically absent.

The presence of hydrogen, carbon monoxide, and carbon dioxide in city gas can cause the following problems:

First, carbon monoxide is highly toxic, and a gas leak poses a risk of poisoning. Its permissible concentration is 200 ppm, and from a safety standpoint, it is desirable to keep the concentration in fuel gas below this level. Even considering dilution by air, it should be kept below 1000 ppm.

Next, carbon dioxide is not only non-flammable, but it also inhibits combustion. Therefore, if it is mixed into fuel gas at high concentrations, it can not only reduce the efficiency of gas transport in pipelines due to a decrease in the calorific value of the fuel gas, but also potentially cause a decrease in the efficiency of combustion equipment.

Finally, although hydrogen is a fuel gas, its calorific value per unit volume is only about one-third that of methane, the main component of city gas. Therefore, when hydrogen is mixed with a fuel gas primarily composed of methane, the calorific value per unit volume decreases. Furthermore, because hydrogen burns quickly, it is known to have a significant impact on combustion equipment.

As described above, hydrogen, carbon monoxide, and carbon dioxide, when mixed with city gas, have various effects at each stage of gas supply and consumption. Therefore, it is common practice to impose restrictions on the concentrations of hydrogen, carbon monoxide, and carbon dioxide in the quality standards for gas accepted into the city gas pipeline network.

In pipeline networks with fuel refueling stations for natural gas vehicles, there are known examples where the upper limit for hydrogen concentration is set at 2% by volume (Non-Patent Document 1). Furthermore, there are known examples where the hydrogen concentration is 4% or less by volume, the carbon dioxide concentration is 0.5% or less by volume, and the carbon monoxide concentration is 0.05% or less by volume (Non-Patent Document 2), as well as examples where the combined concentration of methane and ethane is 93% or more by volume, and the combined concentration of components other than hydrocarbons is 4% or less by volume (Non-Patent Document 3).

The production of methane, the main component gas, by steam reforming of ethanol is thought to proceed through the following reaction. Specifically, hydrogen and carbon monoxide are produced by the steam reforming reaction of ethanol (Equation 1). Some of the carbon monoxide is converted to hydrogen and carbon dioxide by the CO shift reaction (Equation 2), and the remainder reacts with hydrogen to produce methane (Equation 3).
_C2H5OH + _H2O → _2CO + _4H2 (Formula 1)
0.5CO+0.5H ₂ O → 0.5CO ₂ +0.5H ₂ (Formula 2)
1.5CO+4.5H ₂ → 1.5CH ₄ +1.5H ₂ O (Formula 3)
The overall reaction equation is as shown in Equation 4, in which 1 mole of ethanol produces 1.5 moles of methane and 0.5 moles of carbon dioxide.
C ₂ H ₅ OH → 1.5CH ₄ +0.5CO ₂ (Formula 4)

In other words, if ethanol is simply subjected to a methane reaction in the presence of water vapor, even under ideal conditions, the resulting gas will be a mixture of 75% methane and 25% carbon dioxide. Unless the carbon dioxide is removed by some means, it cannot be used as a raw material for city gas.

Methods for removing carbon dioxide include the hot potassium carbonate absorption method and the amine absorption method. These carbon dioxide removal systems have been adopted in alternative natural gas production plants using naphtha and other raw materials. However, all of these carbon dioxide removal methods have high equipment and operating costs, which significantly reduces the economic viability of fuel gas production.

Patent Document 4 discloses a method for producing hydrogen from ethanol, which includes the step of passing the ethanol through a nickel-based catalyst at 300°C to 600°C at an _H2 /ethanol molar ratio of 0.1 to 1.0 and a water vapor/ethanol molar ratio of 1.0 to 10.0 to convert it into a flow containing methane, CO, _CO2 , and _H2 .

This document describes the results of ethanol conversion on a commercially available methanation catalyst at 400°C with a water vapor/ethanol molar ratio of 2.8 and a hydrogen/ethanol molar ratio of 250, achieving a conversion rate of 90.1% and a methane selectivity of 98.8%. It also shows that in ethanol conversion on a commercially available methanation catalyst at 400°C with a water vapor/carbon molar ratio of 3, the absence of hydrogen resulted in a rapid increase in pressure loss in the catalyst layer due to carbon deposition, whereas no increase in pressure loss was observed with a hydrogen/ethanol molar ratio of 0.6. However, in both cases, the resulting gas is primarily hydrogen-based and contains methane, making it unsuitable for use as a raw material for city gas.

Patent Document 5 describes a method for producing a hydrogen-enriched product by selectively reforming ethanol in the presence of a catalyst. This method involves passing ethanol at a molar ratio of 0.2 to 1 to hydrogen through a dehydration/hydrogenation reactor to dehydrate/hydrogenate the ethanol to produce ethane, adiabatically reforming the resulting ethane to create a methane-containing flow, steam reforming the methane-containing flow to produce a mixture containing hydrogen and carbon monoxide, and then performing a water-gas shift reaction to obtain the hydrogen-enriched product.

However, this document only shows the results of the conversion of ethanol to ethane, and does not provide any specific information about the methane concentration in the methane-containing flow.

Another problem when obtaining methane as the main component gas by steam reforming of ethanol is the carbon deposition problem described in Patent Documents 3 and 4. Patent Document 3 shows that significant carbon deposition occurred when a nickel catalyst was used in steam reforming with a water/ethanol ratio of 0.78 (by weight). Patent Document 4 shows that carbon deposition occurred in a short time when no hydrogen was added during ethanol conversion on a commercially available methanation catalyst at 400°C with a water vapor/carbon molar ratio of 3. The water/ethanol molar ratio is 2 in Patent Document 3 and 6 in Patent Document 4. To avoid carbon deposition in steam reforming of ethanol, it is necessary to add a large amount of water vapor. The reaction that produces 1.5 moles of methane and 0.5 moles of carbon dioxide from 1 mole of ethanol (Equation 4) is an exothermic reaction, but the heat generated is only 74 kJ (at 25°C), and even if this is used as the heat of vaporization of water, only about 1.7 moles of water vapor can be produced. Therefore, obtaining methane as the main component gas stably through ethanol steam reforming requires an external supply of steam, which is also economically disadvantageous.

Japanese Unexamined Patent Publication No. 52-52902 Japanese Unexamined Patent Publication No. 55-144093 Japanese Patent Publication No. 2009-227588 Japanese Patent Publication No. 2013-540674 Japanese Patent Publication No. 2006-82996

E.I.Koytsoumpa and S.Karellas, Renewable and Sustainable Energy Reviews, Volume 94, 2018, p.536 Biogas Purchase Guidelines, Osaka Gas Co., Ltd., 2008. Biogas Purchase Guidelines, Tokyo Gas Co., Ltd., 2008.

The problem that this invention aims to solve, in view of the above issues, is to provide an economically advantageous method for producing a high-calorific value fuel gas with a high methane concentration that can be used directly as a raw material for city gas, without using expensive carbon dioxide separation equipment, when obtaining a high-calorific value fuel gas with a methane main component that can be used as city gas by the methanation reaction of ethanol, and without carbon precipitation occurring easily.

The characteristic configuration of the method for producing high-calorific value fuel gas according to the present invention includes: a step of preparing a raw material gas by adding hydrogen and water vapor to ethanol such that the molar ratio of hydrogen/ethanol is 2.0 or more and 2.5 or less, and the molar ratio of water vapor/ethanol is 1.2 or more and 4 or less; a water vapor reforming step of contacting the raw material gas with a catalyst containing ruthenium or nickel at 300°C or more and 700°C or less; and a methane step of contacting the gas obtained in the water vapor reforming step and cooled with a catalyst containing ruthenium or nickel under conditions such that the catalyst outlet temperature is 230°C or more and 330°C or less , wherein the catalyst outlet temperature of the water vapor reforming step is higher than the catalyst outlet temperature of the methane step .

According to this characteristic configuration, the large amount of carbon dioxide produced as a by-product in conventional fuel gas production methods that combine ethanol steam reforming and methane reactions is converted to methane through reaction with hydrogen. Therefore, the large amount of carbon dioxide remaining in the resulting fuel gas can be avoided, and a high-calorific value fuel gas usable as a raw material for city gas can be obtained without the need for expensive carbon dioxide removal equipment. Furthermore, carbon deposition is less likely to occur.

A further characteristic feature of the method for producing high-calorific value fuel gas according to the present invention is that the steam reforming step is carried out under substantially adiabatic conditions with a catalyst inlet temperature of 300°C or higher, and the gas obtained in the steam reforming step is cooled to 230°C to 300°C before being supplied to a heat exchange reactor for the methane process.

According to this characteristic configuration, the steam reforming process is carried out under substantially adiabatic conditions with a catalyst inlet temperature of 300°C or higher. Therefore, the heat generated during the steam reforming process raises the temperature of the steam reforming catalyst layer, maintaining it at a temperature suitable for the steam reforming reaction, thus allowing for sufficient ethanol reforming.

A further characteristic feature of the method for producing high-calorific value fuel gas according to the present invention is that, after adding an amount of oxygen such that the molar ratio to hydrogen in the fuel gas is 0.24 to 0.45 to the methane-based fuel gas obtained in the methane step, the oxygen is reacted with the hydrogen in the fuel gas in the presence of a selective oxidation catalyst capable of selectively oxidizing hydrogen.

Conventional fuel gas production methods, which combine ethanol steam reforming and methane reactions, have the problem of requiring expensive carbon dioxide separation equipment because high concentrations of carbon dioxide remain in the resulting fuel gas. However, due to the high concentration of carbon dioxide, the residual hydrogen concentration tends to be low from the perspective of chemical equilibrium in the methane reaction. In contrast, the high calorific value fuel gas production method of the present invention may result in a higher residual hydrogen concentration. By adopting this characteristic configuration, the residual hydrogen concentration in the fuel gas obtained in the methane step of the high calorific value fuel gas production method of the present invention can be effectively reduced, thereby obtaining a fuel gas with a high methane concentration.

A further characteristic feature of the method for producing high-calorific value fuel gas according to the present invention is that it further includes a dehydration hydrogenation step in which, after adding an amount of ethanol such that the molar ratio to hydrogen in the fuel gas is 0.45 to 0.9 to the fuel gas obtained in the methane step, the ethanol and the hydrogen in the fuel gas are reacted in the presence of a dehydration hydrogenation catalyst to obtain ethane.

By adopting this characteristic configuration, the concentration of hydrogen remaining in the fuel gas obtained in the methane stage of the present invention's method for producing high-calorific value fuel gas can be effectively reduced. Furthermore, by including ethane in addition to methane, it is possible to obtain a high-calorific value fuel gas with a calorific value close to that of typical city gas.

A further characteristic feature of the method for producing high-calorific value fuel gas according to the present invention is that the high-calorific value fuel gas contains methane or methane and ethane, and the proportion of methane or methane and ethane is 93% or more on a volume basis after dehydration.

This characteristic configuration makes it possible to produce high-calorific value fuel gas containing 93% or more methane by volume after dehydration, or high-calorific value fuel gas in which the total amount of methane and ethane is 93% or more by volume after dehydration.

This is a block flow diagram illustrating the method for producing high-calorific value fuel gas according to the present invention. This is a block flow diagram showing another embodiment of the method for producing high-calorific value fuel gas according to the present invention. This is a block flow diagram showing another embodiment of the method for producing high-calorific value fuel gas according to the present invention. This figure shows the relationship between temperature and carbon activity at the water vapor/ethanol ratio ( _H₂O /EtOH) in Examples 1-4 and Comparative Example 1. This figure shows the relationship between temperature and carbon activity at the water vapor/ethanol ratio ( _H₂O /EtOH) of Comparative Examples 2 to 4. This is a process flow diagram showing another embodiment of the method for producing high-calorific value fuel gas according to the present invention. This is a process flow diagram showing another embodiment of the method for producing high-calorific value fuel gas according to the present invention.

[Embodiment]
The following describes embodiments of the method for producing high-calorific value fuel gas according to the present invention. Figure 1 is a block flow diagram showing the method for producing high-calorific value fuel gas according to the present invention. The method for producing high-calorific value fuel gas according to this embodiment includes a step of preparing a raw material gas by adding hydrogen and water vapor to ethanol such that the molar ratio of hydrogen/ethanol is 2.0 or more and 2.5 or less, and the molar ratio of water vapor/ethanol is 1.2 or more and 4 or less (raw material gas preparation section 1); a water vapor reforming step (water vapor reforming reaction section 2) in which the raw material gas is brought into contact with a catalyst containing ruthenium or nickel at 300°C or more and 700°C or less; and a methane reaction step (methanization reaction section 3) in which the gas obtained in the water vapor reforming step is brought into contact with a catalyst containing ruthenium or nickel under conditions such that the catalyst outlet temperature is 230°C or more and 330°C or less. According to the method for producing high-calorific value fuel gas with such a configuration, a high-calorific value fuel gas can be obtained in which the proportion of methane is 93% or more on a volume basis after dehydration.

The ethanol used in the process of preparing the raw material gas does not necessarily have to be produced by fermentation, but from the perspective of producing a fuel gas that can be considered carbon neutral, bioethanol produced from raw materials such as sugarcane or corn is preferable.

Ethanol typically contains small amounts of water, and may also contain trace amounts of organic acids, aldehydes, thiols, and other sulfur compounds. Of these, water may be present at a concentration of approximately 20% to 50% by mass relative to the ethanol. If the raw material ethanol contains water, the amount of added water vapor should be adjusted accordingly so that the molar ratio of water vapor to ethanol supplied to the raw material gas preparation process is between 1.2 and 4.

While trace amounts of organic acids and aldehydes may cause carbon precipitation and other problems in the subsequent steam reforming process if present in excessively large quantities, this is usually not a concern.

Sulfur compounds such as thiols can cause sulfur poisoning in the subsequent steam reforming process, leading to a decrease in activity. Therefore, if a large amount of sulfur compounds are present, it is desirable to perform desulfurization treatment beforehand.

The hydrogen used in the process of preparing the raw material gas can be produced by any method, but hydrogen produced using water electrolysis devices such as alkaline water electrolysis devices, polymer electrolyte water electrolysis devices, and solid oxide water electrolysis devices is preferred because it typically does not contain impurities that would interfere with the steam reforming process.

There are no particular restrictions on the method or order of mixing hydrogen and water vapor with ethanol.

The hydrogen/ethanol molar ratio should be between 2.0 and 2.5. A ratio of 2.05 or higher facilitates a reduction in carbon dioxide concentration in the fuel gas, while a ratio of 2.10 or lower facilitates a reduction in hydrogen concentration. Therefore, a hydrogen/ethanol molar ratio of 2.05 to 2.10 is preferable. However, if a dehydration hydrogenation step using the dehydration hydrogenation reaction of ethanol is included, the hydrogen/ethanol molar ratio may be between 2.2 and 2.5, in which case it is easier to reduce the carbon dioxide concentration in the fuel gas.

The molar ratio of water vapor to ethanol should be between 1.2 and 4.

The reaction that produces 2 moles of methane and 1 mole of water from 1 mole of ethanol and 2 moles of hydrogen generates 156 kJ of heat (at 25°C). If this is entirely used as the heat of vaporization of water, approximately 3.5 moles of water vapor can be produced.

Setting the water vapor/ethanol molar ratio to 1.5 or higher makes it easier to avoid carbon deposition, while setting it to 3 or lower reduces the amount of water vapor required, making it more economically advantageous. Therefore, a water vapor/ethanol molar ratio of approximately 1.5 to 3 is preferable. Furthermore, to suppress the outlet temperature of the water vapor reforming catalyst, an amount of methane resulting in a methane/ethanol molar ratio of approximately 0.5 to 1.5 may be added. In this case, to avoid carbon deposition, it is preferable to set the water vapor/ethanol molar ratio higher than the above, preferably between 2.5 and 4.

As described later, the steam reforming process is an exothermic reaction, and the steam reforming catalyst is exposed to high temperatures, which can pose a durability issue. In such cases, the temperature rise in the steam reforming process can be mitigated by mixing a portion of the outlet gas from the steam reforming process with the raw material gas, thereby diluting the raw material gas. The raw material gas adjustment unit 1 may be equipped with a recycling compressor for recycling a portion of the outlet gas from the steam reforming reaction unit 2, and a flow rate adjustment means for adjusting the amount of recycled gas.

The raw material gas is temperature-adjusted to a range suitable for the subsequent steam reforming process before being introduced into the steam reforming process.

The catalyst used in the steam reforming process is a catalyst in which ruthenium or nickel is supported on an inorganic oxide support such as alumina. The steam reforming process may be carried out in an adiabatic reactor, or in a heat exchange reactor while maintaining a nearly constant catalyst layer temperature.

It is acceptable for a portion of the catalyst layer to be below 300°C. Therefore, the reaction may be initiated by introducing the raw material gas with the catalyst layer inlet temperature below 300°C, for example, around 250°C. However, to ensure the ethanol steam reforming reaction proceeds at a sufficient rate, at least a portion of the catalyst layer must be 300°C or higher. Furthermore, to completely convert ethanol with a small amount of catalyst, it is preferable that at least a portion of the catalyst layer be 400°C or higher. If the catalyst layer inlet temperature is too low, the reaction will proceed significantly slower, requiring an extremely large amount of catalyst, which may negatively impact economic efficiency.

In the steam reforming process, the endothermic reaction of ethanol steam reforming and the exothermic reaction of methane formation occur simultaneously, resulting in an overall exothermic reaction. When a reactor without a special heat removal mechanism (adiabatic reactor) is used, and the raw material gas is introduced at a catalyst layer inlet temperature of around 300°C, and the reaction is carried out under essentially adiabatic conditions, the catalyst layer outlet temperature is typically around 500°C to 700°C.
The steam reforming reaction of ethanol begins at around 300°C and proceeds rapidly above 400°C; therefore, when carried out in an adiabatic reactor, the ethanol reforming reaction is more likely to proceed completely.

Using a heat-exchange reactor, which incorporates a heat exchange function, allows the reaction to proceed while removing the heat generated by the reaction. This suppresses the rise in catalyst layer temperature, which is advantageous in terms of catalyst durability. When the reactor is properly designed and operated, the heat generated by the reaction can be removed, and the catalyst layer inlet and outlet temperatures can be nearly identical, allowing for isothermal reaction. However, it is important to note that even in a heat-exchange reactor, there is a temperature distribution in the catalyst layer, and in areas where exothermic reactions are occurring, the temperature may rise to a level close to that of an adiabatic reaction outlet.

The catalyst used in the methanation process is a catalyst in which ruthenium or nickel is supported on an inorganic oxide support such as alumina. The methanation process may be carried out in an adiabatic reactor, or in a heat exchange reactor while maintaining a nearly constant catalyst layer temperature.

In the methanation process, since the methanation reaction is an exothermic reaction, if the reaction is carried out using an adiabatic reactor, the reaction gas will flow out at a temperature higher than the catalyst inlet temperature. Because the methanation reaction proceeds more towards methane production at lower equilibrium temperatures, lowering the catalyst outlet temperature is advantageous from the viewpoint of obtaining a fuel gas with high methane purity. Therefore, the reaction is carried out under conditions where the catalyst outlet temperature is 330°C or lower. On the other hand, at low temperatures, the activity of the methanation catalyst decreases, and it becomes difficult to carry out the reaction below 230°C. Therefore, the catalyst outlet temperature should be 230°C or higher and 330°C or lower, more preferably 230°C or higher and 280°C or lower.

When a methane reactor is constructed using an adiabatic reactor, multiple reactors are provided, and coolers are installed between the reactors. The inlet temperature of each reaction stage is set to 230°C to 330°C, more preferably 230°C to 280°C, so that the catalyst outlet temperature in the final stage reactor is 230°C to 330°C.

The present invention's method for producing high-calorific value fuel gas may also include a configuration as shown in Figure 2, further comprising a selective oxidation step (selective oxidation reaction section 4a) in which, after adding an amount of oxygen such that the molar ratio to hydrogen in the fuel gas is 0.24 to 0.45 to the methane-based fuel gas obtained in the methane step, the oxygen reacts with the hydrogen in the fuel gas in the presence of a selective oxidation catalyst capable of selectively oxidizing hydrogen. Even with such a configuration, a high-calorific value fuel gas can be obtained in which methane accounts for 93% or more of the volume after dehydration.

The oxygen used in the selective oxidation reaction step can be produced by any method, as long as it has sufficient purity and properties to prevent interference with the selective oxidation reaction. For example, when hydrogen is obtained by the electrolysis of water, an amount of oxygen with a molar ratio (oxygen/hydrogen) of 0.5 is inevitably produced, and a portion of this oxygen can be used. Since the oxygen obtained by the electrolysis of water is usually of high purity, this method is economically advantageous.

The reaction in selective oxidation reaction section 4a proceeds as follows.
H ₂ + 0.5O ₂ → H ₂ O (Formula 5)

If the amount of oxygen added is too low, the reduction in hydrogen concentration in the fuel gas will be insufficient. On the other hand, if the amount of oxygen added is too high, oxygen may remain in the fuel gas, and a rapid oxidation reaction may occur on the selective oxidation catalyst, potentially leading to catalyst degradation. Therefore, it is preferable to set the molar ratio of oxygen to hydrogen (oxygen/hydrogen) in the fuel gas obtained in the methane reaction step to 0.24 or higher and 0.45 or lower.

Preferably, a selective oxidation catalyst is active in the oxygen-hydrogen reaction (Equation 5) but substantially inactive in methane-consuming reactions such as the oxidation of methane by oxygen or the steam reforming of methane. A catalyst that shows high activity in methane oxidation not only fails to effectively reduce hydrogen concentration but also reduces the amount of methane obtained as fuel gas, potentially lowering the efficiency of fuel gas production. Using a catalyst active in the steam reforming of methane may result in insufficient reduction of hydrogen concentration in the fuel gas, as new hydrogen is generated by the steam reforming reaction even if hydrogen is reduced by the reaction with oxygen. An example of a catalyst exhibiting such reaction selectivity is a catalyst in which at least one of palladium or platinum is supported on an inorganic oxide support.

When selectively oxidizing hydrogen in methane gas using a selective oxidation catalyst, the reaction temperature is preferably between 100°C and 400°C. Selective oxidation catalysts generally exhibit good activity at temperatures above 100°C, so a reaction temperature above 100°C facilitates the reaction between oxygen and hydrogen. Furthermore, a reaction temperature below 400°C tends to suppress the steam reforming reaction of methane. A reaction temperature of 150°C to 300°C is more preferable.

The type of reactor used in the selective oxidation process is not particularly limited and may include, for example, a fixed-bed insulated reactor, a fixed-bed insulated reactor with a recycling line, a heat exchange reactor, etc.

The present invention's method for producing a high-calorific value fuel gas may further include a dehydration hydrogenation step (dehydration hydrogenation reaction section 4b), as shown in Figure 3. This step involves adding ethanol to a fuel gas obtained in a methanation step, where methane is the main component and hydrogen is present, in an amount such that the molar ratio to hydrogen in the fuel gas is 0.45 to 0.9. The ethanol is then reacted with the hydrogen in the fuel gas in the presence of a dehydration hydrogenation catalyst to obtain ethane. With this configuration, a high-calorific value fuel gas can be obtained in which methane and ethane account for 93% or more of the volume after dehydration.

The ethanol used in the dehydration and hydrogenation reaction step may be the same as the ethanol used in the preparation of the raw material gas.

The reaction in the dehydration and hydrogenation reaction section 4b proceeds as follows: In the dehydration and hydrogenation reaction section 4b, the dehydration reaction of ethanol (Equation 6) and the hydrogenation reaction of ethylene (Equation 7) proceed.
C ₂ H ₅ OH → C ₂ H ₄ +H ₂ O (Formula 6)
C ₂ H ₄ + H ₂ → C ₂ H ₆ (Formula 7)

If the amount of ethanol added is insufficient, the reduction in hydrogen concentration in the fuel gas will be inadequate.
On the other hand, if the amount of ethanol added is too large, the hydrogen concentration becomes extremely low, which can prevent the hydrogenation reaction of ethylene (Equation 7) from proceeding sufficiently, resulting in residual ethylene in the fuel gas. Furthermore, the high ethylene concentration during the reaction can lead to polymerization of ethylene on the catalyst, causing catalyst degradation due to carbon deposition. Therefore, the amount of ethanol added should be such that the molar ratio of ethanol to hydrogen in the fuel gas obtained in the methane process (ethanol)/(hydrogen) is between 0.45 and 0.90.

The dehydration hydrogenation catalyst used in the dehydration hydrogenation reaction step is preferably active in the dehydration reaction of ethanol (Equation 6) and the hydrogenation reaction of olefins (Equation 7), while being substantially inactive in the steam reforming reactions of methane and ethane. If a catalyst active in the steam reforming reactions of methane and ethane is used, even if hydrogen is reduced in the hydrogenation reaction of ethylene, new hydrogen may be generated by the steam reforming reactions of methane and ethane, making it impossible to reduce the hydrogen concentration in the fuel gas. An example of a catalyst exhibiting such reaction selectivity is a catalyst in which at least one of palladium or platinum is supported on a solid acid catalyst.

The reaction temperature when converting ethanol to ethane by reaction with hydrogen using a dehydration hydrogenation catalyst is preferably between 200°C and 400°C. Since dehydration hydrogenation catalysts generally exhibit good activity at temperatures above 200°C, a reaction temperature of 200°C or higher facilitates the reaction between ethanol and hydrogen (Equations 6 and 7). Furthermore, a reaction temperature below 400°C tends to suppress the steam reforming reactions of methane and ethane. A reaction temperature of 250°C to 350°C is more preferable.

The type of reactor used in the dehydration and hydrogenation reaction process is not particularly limited and may include, for example, a fixed-bed insulated reactor, a fixed-bed insulated reactor with a recycling line, a heat exchange reactor, etc.

The pressure in the steam reforming process should be 0.5 MPa (absolute pressure, the same applies hereafter) or higher to easily obtain a sufficient reaction rate, and 5 MPa or lower to reduce equipment costs and be economically advantageous. Therefore, a pressure of 0.5 MPa to 5 MPa is preferable, and a pressure of 0.5 MPa to 3 MPa is even more preferable.

The pressure in the methanation process should be 0.5 MPa or higher, as this makes it easier to obtain a sufficient reaction rate and ensures that methanation proceeds smoothly in equilibrium. A pressure of 5 MPa or lower is more economically advantageous because it reduces equipment costs. Therefore, a pressure of 0.5 MPa to 5 MPa is preferable, and a pressure of 0.5 MPa to 3 MPa is even more preferable.

The pressure in the selective oxidation and dehydration/hydrogenation steps should be 0.5 MPa or higher to easily obtain a sufficient reaction rate, and 5 MPa or lower to reduce equipment costs and be economically advantageous. Therefore, a pressure of 0.5 MPa to 5 MPa is preferable, and a pressure of 0.5 MPa to 3 MPa is more preferable.

The steam reforming, methane, selective oxidation, and dehydration/hydrogenation processes are all preferably conducted at a pressure of 0.5 MPa to 5 MPa, and more preferably at 0.5 MPa to 3 MPa. Therefore, it is not usually necessary to change the pressure between each process, but pressurization or depressurization steps may be added as needed. Normally, the reaction pressure decreases slightly in accordance with the pressure loss in each process, but this does not pose any particular problem.

[Examples and Comparative Examples]
The following is an example of a calculation based on process calculations. Pressure loss and heat loss in each piece of equipment and piping are not considered.

[Example 1]
Ethanol, hydrogen, and water are supplied to the raw material gas adjustment unit 1 in a molar ratio of 1:2:1.2 and at a pressure of 0.7 MPa (absolute pressure). In the raw material gas adjustment unit 1, these are mixed and heated to 300°C to adjust the raw material gas, which is then sent to the steam reforming reaction unit 2.

In the steam reforming reaction section 2, the raw material gas comes into contact with the steam reforming catalyst under adiabatic conditions, and the steam reforming reaction of ethanol, the CO shift reaction, and the methane reaction proceed until a chemical equilibrium state is reached. The generated gas is cooled to 250°C and supplied to the methane reaction section 3.

In the methanation reaction section 3, the CO shift reaction and methanation reaction proceed under isothermal conditions of 250°C, reaching a state of chemical equilibrium and yielding a fuel gas mainly composed of methane and water vapor. Cooling this fuel gas and separating the water through condensation yields a fuel gas mainly composed of methane, with small amounts of hydrogen and carbon dioxide.

Table 1 shows the inlet and outlet temperatures in each reaction chamber, as well as the composition of the outlet gas (based on volume after dehydration) and the carbon activity. Here, the carbon activity is calculated as _KP /( _PCO2 / _PCO2 ) using the partial pressures of carbon monoxide and carbon dioxide in the outlet gas ( _PCO2 , _PCO2 ) and the equilibrium constant ( _KP ) of the thermodynamically calculated carbon monoxide disproportionation reaction (Equation ⁸ ). When this value is greater than 1, the carbon deposition reaction proceeds in chemical equilibrium.
2CO → C (solid) + CO ₂ (Formula 8)

The outlet temperature of the steam reforming reaction section 2 is 619°C. The gas composition at the outlet of the steam reforming reaction section 2 is 39.44% methane and 47.48% hydrogen. The methane concentration is less than 40%, resulting in a gas containing a high concentration of hydrogen.

The gas composition at the outlet of the methanation reaction section 3 is 96.20% methane, 3.04% hydrogen, and 0.76% carbon dioxide, resulting in a fuel gas (high calorific value fuel gas) with a methane purity high enough to be used as a raw material for city gas.

At both the outlet of the steam reforming reaction section 2 and the outlet of the methane reaction section 3, the carbon activity is less than 1, thus avoiding carbon deposition in terms of chemical equilibrium.

In the above, it was assumed that the steam reforming reaction occurred under adiabatic conditions and the methane reaction under isothermal conditions. However, even in adiabatic reactors, a certain amount of heat is dissipated from the reactor surface, and localized temperature increases can occur even in heat exchange reactors. From this perspective, it is desirable that the carbon activity be less than 1 when the steam reforming, CO shift, and methane reactions of ethanol reach equilibrium, under conditions where ethanol, hydrogen, and water are in a molar ratio of 1:2:1.2, at a pressure of 0.7 MPa, and the outlet temperature in the adiabatic reaction is 619°C or lower.

Figure 4 (water vapor/ethanol = 1.2) shows the carbon activity at chemical equilibrium at various temperatures. Since the carbon activity is below 1 across the entire temperature range from 250°C to 619°C, carbon precipitation is avoided in chemical equilibrium under adiabatic, isothermal, and intermediate conditions.

[Example 2]
The calculations were performed in the same manner as in Example 1, except that the molar ratio of ethanol, hydrogen, and water supplied to the raw material gas adjustment unit 1 was set to ethanol:hydrogen:water = 1:2:1.5.

Table 2 shows the inlet and outlet temperatures, as well as the composition (by volume after dehydration) and carbon activity of the outlet gas in each reaction chamber.

The outlet temperature of the steam reforming reaction section 2 is 608°C. The gas composition at the outlet of the steam reforming reaction section 2 is 39.14% methane and 47.86% hydrogen. Due to the increased amount of added steam, the methane reaction, which produces water as a product, becomes less likely to proceed in equilibrium, resulting in a slight decrease in methane concentration compared to Example 1.

The gas composition at the outlet of the methanation reaction section 3 is 95.90% methane, 3.28% hydrogen, and 0.82% carbon dioxide. While the methane concentration decreases slightly, it becomes a fuel gas (high-calorific value fuel gas) with a methane purity high enough to be used as a raw material for city gas.

At both the outlet of the steam reforming reaction section 2 and the outlet of the methanation reaction section 3, the carbon activity is below 1, which is lower than in Example 1. This is thought to be because increasing the amount of steam added makes carbon precipitation less likely. As shown in Figure 4 (steam/ethanol = 1.5), the carbon activity is below 1 throughout the entire temperature range from 250°C to 608°C, and in equilibrium, carbon precipitation does not occur in the entire temperature range expected for steam reforming and methanation reactions.

[Example 3]
The calculations were performed in the same manner as in Example 1, except that the molar ratio of ethanol, hydrogen, and water supplied to the raw material gas adjustment unit 1 was set to ethanol:hydrogen:water = 1:2:2.

Table 3 shows the inlet and outlet temperatures, as well as the composition (by volume after dehydration) and carbon activity of the outlet gas in each reaction chamber.

The gas composition at the outlet of the methanation reaction section 3 is 95.42% methane, 3.66% hydrogen, and 0.92% carbon dioxide. While the methane concentration is slightly lower than in Example 2, it is a fuel gas (high-calorific value fuel gas) with a methane purity high enough to be used as a raw material for city gas.

At both the outlet of the steam reforming reaction section 2 and the outlet of the methanation reaction section 3, the carbon activity is well below 1. As shown in Figure 4 (steam/ethanol = 2), the carbon activity remains below 1 throughout the entire temperature range from 250°C to 591°C, and in equilibrium, no carbon precipitation occurs across the entire temperature range expected for steam reforming and methanation reactions.

[Example 4]
The calculations were performed in the same manner as in Example 1, except that the molar ratio of ethanol, hydrogen, and water supplied to the raw material gas adjustment unit 1 was set to ethanol:hydrogen:water = 1:2:3.

Table 4 shows the inlet and outlet temperatures, as well as the composition (by volume after dehydration) and carbon activity of the outlet gas in each reaction chamber.

The outlet temperature of the steam reforming reaction section 2 is 562°C. This is a decrease of more than 50°C compared to Example 1. This is because the increased amount of steam added mitigates the temperature rise, which is advantageous from the standpoint of the durability of the steam reforming catalyst.

The gas composition at the outlet of the methanation reaction section 3 is 94.51% methane, 4.39% hydrogen, and 1.10% carbon dioxide. While the methane concentration is slightly lower than in Example 3, it is a fuel gas (high calorific value fuel gas) with a methane purity high enough to be used as a raw material for city gas.

At both the outlet of the steam reforming reaction section 2 and the outlet of the methane reaction section 3, the carbon activity was well below 1. Furthermore, as shown in Figure 4 (steam/ethanol = 3), the carbon activity remained below 1 throughout the entire temperature range from 250°C to 562°C, indicating that, in equilibrium, no carbon precipitation occurs across the entire temperature range expected for steam reforming and methane reactions.

[Comparative Example 1]
The calculations were performed in the same manner as in Example 1, except that the molar ratio of ethanol, hydrogen, and water supplied to the raw material gas adjustment unit 1 was set to ethanol:hydrogen:water = 1:2:1.

Table 5 shows the inlet and outlet temperatures, as well as the composition (by volume after dehydration) and carbon activity of the outlet gas in each reaction chamber.

The outlet temperature of the steam reforming reaction section 2 reached 628°C. Furthermore, the carbon activity at the outlet of the steam reforming reaction section 2 exceeded 1. Compared to the example, the lower amount of steam added resulted in a higher outlet temperature and a carbon activity exceeding 1, raising concerns about deterioration of the steam reforming catalyst due to thermal degradation and carbon deposition.

As shown in Figure 4 (water vapor/ethanol = 1), carbon deposition does not occur in equilibrium below 580°C. Therefore, it may be possible to avoid carbon deposition by proceeding with the reaction while removing heat. However, the temperature may locally reach a level close to the exit temperature of the adiabatic reaction, and controlling this is somewhat difficult.

[Comparative Example 2]
The calculations were performed in the same manner as in Example 1, except that the molar ratio of ethanol and water supplied to the raw material gas adjustment unit 1 was set to ethanol:water = 1:2, and no hydrogen was added.

Table 6 shows the inlet and outlet temperatures, as well as the composition (by volume after dehydration) and carbon activity of the outlet gas in each reaction chamber.

The outlet temperature of the steam reforming reaction section 2 is 518°C. The gas composition at the outlet of the steam reforming reaction section 2 is 49.36% methane and 25.09% hydrogen. The methane concentration is higher compared to Example 1, which is thought to be because the lower outlet temperature compared to Example 1 allows the methanation reaction to proceed more easily in equilibrium.

On the other hand, the gas composition at the outlet of the methanation reaction section 3 is 73.87% methane, 1.13% hydrogen, and 25.00% carbon dioxide. Unless the carbon dioxide is removed, it cannot be used as a raw material for city gas.

A further problem is that the carbon activity exceeds 1 at both the outlet of the steam reforming reaction section 2 and the outlet of the methane reaction section 3, raising concerns about degradation of the steam reforming catalyst and methane catalyst due to carbon deposition. Figure 5 (steam/ethanol = 2.0) shows the carbon activity at each temperature when chemical equilibrium is reached. Since the carbon activity exceeds 1 across the entire temperature range below 700°C, carbon deposition proceeds in chemical equilibrium under adiabatic, isothermal, and intermediate conditions.

[Comparative Example 3]
The calculations were performed in the same manner as in Comparative Example 2, except that the molar ratio of ethanol and water supplied to the raw material gas adjustment unit 1 was set to ethanol:water = 1:2.5.

Table 7 shows the inlet and outlet temperatures, as well as the composition (by volume after dehydration) and carbon activity of the outlet gas in each reaction chamber.

The temperature at the outlet of the steam reforming reaction section 2 reached 502°C, and the carbon activity fell below 1.

The gas composition at the outlet of the methanation reaction section 3 is 73.66% methane, 1.33% hydrogen, and 25.00% carbon dioxide. Unless the carbon dioxide is removed, it cannot be used as a raw material for city gas. Furthermore, the carbon activity at the outlet of the methanation reaction section 3 exceeds 1, raising concerns about degradation of the methanation catalyst due to carbon deposition. Figure 5 (water vapor/ethanol = 2.5) shows the carbon activity at chemical equilibrium at various temperatures. Below 340°C, the carbon activity exceeds 1, indicating that if the methanation reaction is carried out at low temperatures to promote methanation, carbon deposition will proceed in chemical equilibrium.

[Comparative Example 4]
The calculations were performed in the same manner as in Comparative Example 2, except that the molar ratio of ethanol and water supplied to the raw material gas adjustment unit 1 was set to ethanol:water = 1:3.

Table 8 shows the inlet and outlet temperatures, as well as the composition (by volume after dehydration) and carbon activity of the outlet gas in each reaction chamber.

The temperature at the outlet of the steam reforming reaction section 2 reached 488°C.

The gas composition at the outlet of the methanation reaction section 3 is 73.46% methane, 1.54% hydrogen, and 25.00% carbon dioxide. Unless carbon dioxide is removed, it cannot be used as a raw material for city gas. At both the outlet of the steam reforming reaction section 2 and the outlet of the methanation reaction section 3, the carbon activity is below 1, and carbon deposition is avoided in chemical equilibrium. The results of Comparative Examples 2-4 indicate that in simple steam reforming of ethanol without the presence of hydrogen, to avoid carbon deposition in equilibrium, it is necessary to add steam so that the molar ratio of steam to ethanol is 3 or more (the molar ratio to the produced methane is 2 or more). Note that, as shown in Figure 5 (steam/ethanol = 3.0), below 245°C, the carbon activity exceeds 1. Therefore, if the methanation reaction is carried out below 245°C to promote methanation, carbon deposition may proceed in chemical equilibrium.

[Example 5]
An example of a fuel gas production process with a selective oxidation step is shown. The process flow is shown in Figure 6.

The raw material gas adjustment unit 1, where the raw material gas adjustment process takes place, is supplied with ethanol, hydrogen, and water at flow rates of 1 mol/s, 2.08 mol/s, and 2.5 mol/s, respectively, at a temperature of 25°C and a pressure of 0.8 MPa. The raw material gas adjustment unit 1 is equipped with heat exchangers 11-14 and a recycling compressor 15.

Ethanol, hydrogen, and water are heated and mixed, then further mixed with the recycled gas from the outlet of the steam reforming reactor 21, heated to 300°C, and become the raw material gas, which is then supplied to the steam reforming reaction section 2.

The steam reforming reaction unit 2, where the steam reforming process takes place, comprises a steam reforming reactor 21 and a heat exchanger 22. The raw material gas is fed into the steam reforming reactor 21, where the steam reforming reaction of ethanol, the CO shift reaction, and the methane reaction proceed adiabatically. The outlet gas from the steam reforming reactor 21 is cooled to 240°C in the heat exchanger 22, after which one-third is returned to the raw material gas adjustment process and mixed with the raw material gas through the recycling compressor 15. The remainder is fed into the methane reaction unit 3.

The methanation reaction unit 3, where the methanation process takes place, comprises a first methanation reactor 31, a heat exchanger 32, a second methanation reactor 33, a heat exchanger 34, a third methanation reactor 35, and a heat exchanger 36.

The first, second, and third methanation reactors 31, 33, and 35 are all adiabatic reactors. The outlet gases from the first and second methanation reactors 31 and 33 are cooled to 240°C in heat exchangers 32 and 34, respectively, before being sequentially supplied to the next stage reactors. The outlet gas from the third methanation reactor 35 is cooled to 200°C in heat exchanger 36 before being supplied to the selective oxidation process.

The selective oxidation reaction section 4a, where the selective oxidation process takes place, comprises a selective oxidation reactor 41a and a heat exchanger 42a. In the selective oxidation reaction section 4a, 0.02 mol/s of oxygen at 200°C is added to the outlet gas of the third methanation reactor 35. On the selective oxidation catalyst packed in the selective oxidation reactor 41a, the hydrogen contained in the outlet gas of the third methanation reactor 35 reacts with the oxygen to form water vapor, yielding a fuel gas with reduced hydrogen concentration. This fuel gas is cooled in the heat exchanger 42a, and water is separated and dehydrated to obtain a fuel gas mainly composed of methane.

Table 9 shows the temperature and flow rate at key points in the process.

The composition of the resulting fuel gas (by volume after dehydration) is 95.83% methane, 3.72% hydrogen, 0.45% carbon dioxide, and 2 ppm carbon monoxide. This demonstrates that the method of the present invention can produce a fuel gas with high methane purity (high calorific value fuel gas) that can be used as city gas simply by adjusting the calorific value with propane or the like.

Furthermore, the carbon activity at the outlet of each reactor is less than 1, thus avoiding carbon deposition in terms of chemical equilibrium.

[Example 6]
An example of a fuel gas production process that includes a dehydration and hydrogenation step is shown. The process flow is shown in Figure 7.

In the raw material gas preparation process, ethanol, hydrogen, and water are supplied at a temperature of 25°C and a pressure of 0.8 MPa, at flow rates of 1 mol/s, 2.25 mol/s, and 2.5 mol/s, respectively. In the raw material gas preparation process, these are heated to 220°C and mixed, then mixed with the recycled gas from the outlet of the steam reforming reactor 21. This mixture is then heated to 300°C to become the raw material gas, which is then sent to the steam reforming process.

In the steam reforming process, the raw material gas is introduced into the steam reforming reactor 21, where the reaction proceeds adiabatically. After the outlet gas from the steam reforming reactor 21 is cooled to 240°C, one-third of it is returned to the raw material gas adjustment process via the recycling compressor 15, and the remainder is introduced into the first methanation reactor 31. The first and second methanation reactors 31 and 33 are both adiabatic reactors. The outlet gas from the first methanation reactor 31 is cooled to 240°C before being introduced into the second methanation reactor 33. The outlet gas from the second methanation reactor 33 is cooled to 250°C before being introduced into the dehydration and hydrogenation reaction process.

In the dehydration hydrogenation reaction step, 0.25 mol/s of ethanol at 250°C is added to the outlet gas of the second methanation reactor 33 and sent to the dehydration hydrogenation reactor 41b, which is filled with a dehydration hydrogenation catalyst. On the dehydration hydrogenation catalyst, the hydrogen contained in the outlet gas of the second methanation reactor 33 reacts with the ethanol to form ethane, reducing the hydrogen concentration and yielding a fuel gas containing ethane. Cooling this fuel gas yields a fuel gas mainly composed of methane, also containing ethane.

Table 10 shows the temperature and flow rate at key points in the process.

The composition of the resulting fuel gas (by volume after dehydration) is 86.82% methane, 10.91% ethane, 1.82% hydrogen, 0.45% carbon dioxide, 161 ppm ethylene, and 12 ppm carbon monoxide. This demonstrates that the method of the present invention can produce a high-calorific value fuel gas containing methane and ethane that can be used as city gas with only minor adjustments to its calorific value.

Furthermore, the carbon activity at the outlet of each reactor is less than 1, thus avoiding carbon deposition in terms of chemical equilibrium.

The following test examples relate to the steam reforming step in the high-calorific value fuel gas production method of the present invention.

[Test Example 1]
A temperature measuring sheath tube (6 mm outer diameter) was passed through the center of a stainless steel reaction tube (20 mm inner diameter). 12 mL (approximately 9 g) of catalyst, consisting of activated alumina supports (2-4 mm spherical), bearing 2% by mass of ruthenium, was packed between the reaction tube and the sheath tube to form a catalyst layer. Alumina balls were packed to a height of 30 mm both above and below the catalyst layer. This reaction tube was loaded into an electric furnace, and while heating until the temperature of the center of the alumina ball layer above the catalyst layer reached 250°C, a reducing gas mixture of nitrogen gas and 2% hydrogen gas (by volume) was circulated, and the reduction treatment was carried out for 1 hour.

After the reduction treatment described above, the pressure inside the reaction tube was maintained at 0.7 MPa (absolute pressure), and a mixture of ethanol (200 mL/min at standard conditions of 0°C and 1 atm, 15 mm above the top of the catalyst layer), hydrogen (400 mL/min), and water vapor (400 mL/min) was flowed through the catalyst layer from top to bottom while heating the center of the alumina ball layer on the catalyst layer, i.e., 15 mm above the top of the catalyst layer, to reach a temperature of 250°C. The gas at the outlet of the catalyst layer was separated from condensed components (water, ethanol) using an ice-cooled trap. Carbon monoxide, carbon dioxide, hydrogen, and nitrogen in the gas were analyzed using a gas chromatograph (Agilent Micro-GC with TCD detector), and hydrocarbons (methane, ethylene, ethane, propylene, propane, butane) and ethanol were analyzed using a gas chromatograph (Shimadzu GC-2014 with FID detector). Furthermore, the ethanol in the condensed water obtained using the ice-cooled trap was analyzed using a gas chromatograph (Shimadzu GC-2014, with FID detector).

The temperature control point was chosen to be the alumina ball layer above the top of the catalyst layer, rather than inside the catalyst layer itself, because accurately measuring the temperature inside the catalyst layer is difficult due to the heat generated by the reaction. Since the reaction does not proceed within the alumina ball layer, the gas temperature at the catalyst layer inlet is approximately the same as that at the center of the alumina ball layer. In the following, "catalyst inlet temperature" refers to the temperature at the center of the alumina ball layer above the catalyst layer. The temperatures inside the catalyst layer (at three points: top, middle, and bottom) were also measured using thermocouples inserted into the sheath tube.

After the measurement at 250°C was completed, the catalyst inlet temperature was sequentially changed to 300°C, 350°C, and 400°C while the test gas was still flowing, and the catalyst outlet gas was analyzed similarly using gas chromatography.

The ethanol conversion rate in the reaction was calculated using the following formula. The ethanol contained in the catalyst outlet gas includes not only the gas that passed through the ice-cooled trap, but also the ethanol dissolved in the water condensed in the ice-cooled trap.
Ethanol conversion rate [%] = 100 × {1 - (amount of ethanol per unit time contained in the catalyst outlet gas) / (amount of ethanol per unit time supplied to the catalyst layer)}

Table 11 shows the temperatures of the upper, middle, and lower parts of the catalyst layer at each catalyst inlet temperature, as well as the analysis results of the gas composition after condensate separation and the ethanol conversion rate.

At a catalyst inlet temperature of 250°C, the generated gas was primarily methane, containing 88.7% methane, 9.3% hydrogen, and 2.0% carbon dioxide. The ethanol conversion rate was 98.0%, with a small amount of ethanol remaining in the condensate. The maximum temperature within the catalyst layer was 339°C (upper part), and it was 284°C at the catalyst layer outlet. Although the temperature within the catalyst layer rises above the catalyst inlet temperature due to the heat generated by methane production, it can be seen that the temperature decreases within the catalyst layer due to heat dissipation because the catalyst layer is not sufficiently insulated.

At a catalyst inlet temperature of 300°C, the generated gas was primarily methane, containing 87.6% methane. The ethanol conversion rate was 100%, and no ethanol was detected in either the gas or the condensed water. The maximum temperature within the catalyst layer was 371°C. Compared to the case with a catalyst inlet temperature of 250°C, the higher maximum temperature within the catalyst layer suggests that complete ethanol conversion was achieved. Note that in the process calculations for adiabatic steam reforming shown in Table 3, the catalyst layer outlet temperature was 591°C. However, under the conditions of this test example, the high heat dissipation likely resulted in the maximum temperature within the catalyst layer remaining at 371°C. On the other hand, the lower catalyst layer temperature was 319°C, which facilitated the methanation reaction to proceed in equilibrium, resulting in a significantly higher methane concentration in the generated gas compared to Table 3 (38.83%).

When the catalyst inlet temperature was set to 350°C and 400°C, the methane concentration of the generated gas was slightly lower compared to when the catalyst inlet temperature was 300°C, due to the higher catalyst outlet temperature.

[Test Example 2]
The experiment was conducted in the same manner as in Test Example 1, except that half of the hydrogen supplied to the catalyst layer was replaced with nitrogen, resulting in a flow rate of 200 mL/min of ethanol, 200 mL/min of hydrogen, 200 mL/min of nitrogen, and 400 mL/min of water vapor. Nitrogen was added to ensure that the total flow rate of gas supplied to the catalyst per unit time was the same as in Test Example 1.

Table 12 shows the temperatures of the upper, middle, and lower parts of the catalyst layer at each catalyst inlet temperature, as well as the analysis results of the gas composition after condensate separation and the ethanol conversion rate.

At a catalyst inlet temperature of 250°C, the generated gas consisted of 51.6% methane, 5.9% hydrogen, and 8.7% carbon dioxide, with trace amounts (less than 0.02%) of ethane and propane, resulting in a methane-dominant gas. Even after removing nitrogen, a large amount of carbon dioxide remained in the generated gas; this is because, stoichiometrically, there was insufficient hydrogen to convert all the carbon in the ethanol into methane. The ethanol conversion rate was 97.8%, and in addition to condensed water, some ethanol remained in the gas. The maximum temperature within the catalyst layer was 363°C (upper part), and 336°C in the lower part of the catalyst layer. At catalyst inlet temperatures above 300°C, the ethanol conversion rate reached 100%, and the only hydrocarbon in the generated gas was methane.

[Test Example 3]
The test was conducted in the same manner as in Test Example 1, except that the entire amount of hydrogen supplied to the catalyst layer was replaced with nitrogen, resulting in ethanol at 200 mL/min, nitrogen at 400 mL/min, and water vapor at 400 mL/min.

Table 13 shows the temperatures of the upper, middle, and lower parts of the catalyst layer at each catalyst inlet temperature, as well as the analysis results of the gas composition after condensate separation and the ethanol conversion rate.

At a catalyst inlet temperature of 250°C, the ethanol conversion rate was only 11.5%. While the ethanol conversion rate increased with increasing catalyst inlet temperature, it remained at 64.6% even at 400°C. Furthermore, at all temperatures, the generated gas (after excluding nitrogen) was primarily hydrogen, containing carbon monoxide, carbon dioxide, methane, and small amounts (0.02% to 4.0%) of C2-C4 hydrocarbons (ethylene, ethane, propylene, propane, and butane).

[Evaluation of Test Examples 1-3]
In Test Example 1, according to the present invention, hydrogen and water vapor were added to ethanol so that the molar ratio of hydrogen/ethanol was 2.0 and the molar ratio of water vapor/ethanol was 2.0, and the mixture was brought into contact with a ruthenium-containing catalyst. When the catalyst inlet temperature was 250°C or higher, the ethanol was almost completely converted, and a gas mainly composed of methane was obtained. Furthermore, when the mixture was brought into contact with the catalyst under conditions where the catalyst inlet temperature was 300°C or higher and the maximum temperature in the catalyst layer was 370°C or higher, the ethanol could be completely converted into a gas mainly composed of methane.

On the other hand, in Test Example 2, where the hydrogen/ethanol molar ratio was 1.0, the ethanol conversion rate was about the same as in the case of a hydrogen/ethanol molar ratio of 2.0, but a large amount of carbon dioxide remained in the generated gas.

Furthermore, in Test Example 3, where hydrogen was not added, ethanol did not completely convert even at a catalyst inlet temperature of 400°C (catalyst layer internal temperature of 392°C to 398°C). As disclosed in Patent Document 3, a temperature of 400°C or higher is required to obtain a sufficient ethanol conversion rate in a simple ethanol steam reforming reaction without hydrogen addition.

Examples 7 and Comparative Example 5 are examples in which catalysts were packed in two stages at the inlet and outlet ends of a long reaction tube, and gases containing ethanol, hydrogen, and water vapor were brought into contact with the catalysts at different temperatures to carry out a water vapor reforming reaction at the inlet end and a methane reaction at the outlet end, in an attempt to produce a high-calorific value fuel gas.

[Example 7]
A temperature measuring sheath tube (6 mm outer diameter) was passed through the center of a stainless steel reaction tube (20 mm inner diameter). Between the reaction tube and the sheath tube, 12 mL (approximately 9 g) of catalyst, in which 2 mass% of ruthenium was supported on activated alumina supports (2-4 mm spherical), was packed to form a steam reforming catalyst layer. Above the steam reforming catalyst layer, alumina balls were packed to a height of 30 mm. Below the steam reforming catalyst layer, alumina balls were packed to a height of 145 mm to form a cooling zone for cooling the gas emitted from the steam reforming catalyst. Further below that, 43 mL (approximately 33 g) of catalyst, in which 2 mass% of ruthenium was supported on activated alumina supports (2-4 mm spherical), was packed to form a methanation catalyst layer. Thermocouples for temperature measurement were installed inside the sheath tube at the center of the alumina ball layer above the steam reforming catalyst layer, at the upper, middle, and lower parts of the steam reforming catalyst layer, and at the upper, middle, and lower parts of the methane catalyst layer.

The reaction tube was loaded into the electric furnace. The electric furnace consists of three zones: upper, middle, and lower, with the heating output of each zone individually controllable. The reaction tube was loaded into the electric furnace so that the lower end of the steam reforming catalyst layer was at the same height as the lower end of the upper heater, and the upper and lower ends of the lower heater were at the same height as the upper and lower ends of the methane catalyst layer.

The upper heater was controlled to maintain a temperature of approximately 250°C at the center of the alumina ball layer above the steam reforming catalyst layer, and the lower heater was controlled to maintain a temperature of 250°C at the outlet of the methane catalyst layer. A reducing gas, a mixture of nitrogen gas and 2% hydrogen gas (by volume), was then circulated, and the reduction treatment was performed for one hour.

After the reduction treatment described above, the pressure inside the reaction tube was maintained at 0.7 MPa (absolute pressure). The upper heater was controlled to maintain a temperature of 300°C at the center of the alumina ball layer above the steam reforming catalyst layer (steam reforming catalyst inlet temperature), while the lower heater was controlled to maintain a temperature of 270°C at the outlet (bottom) of the methane catalyst layer. A mixture of 200 mL/min of ethanol, 400 mL/min of hydrogen, and 400 mL/min of steam was circulated from top to bottom through the catalyst layer. The catalyst layer outlet gas was analyzed in the same manner as in Test Example 1.

Table 14 shows the steam reforming catalyst inlet temperature, the upper, middle, and lower temperatures of the steam reforming catalyst layer, the upper, middle, and lower temperatures of the methane catalyst layer, and the analysis results of the gas composition after condensate separation, as well as the ethanol conversion rate, at 1, 2, 3, and 4 hours after the start of the reaction.

One hour after the start of the reaction, the steam reforming catalyst inlet temperature was 295°C, the upper steam reforming catalyst temperature was 392°C, and the methanation catalyst outlet temperature (lower part) was 269°C. The gas after condensate separation consisted of 94.2% methane, 5.3% hydrogen, and 0.5% carbon dioxide, resulting in a methane-dominant gas with small amounts of hydrogen and carbon dioxide. The ethanol conversion rate was 100%.

From two hours after the start of the reaction, the steam reforming catalyst inlet temperature was stably controlled at 299–300°C, the upper steam reforming catalyst layer temperature was 397–401°C, and the methanation catalyst outlet temperature (lower part) was 270–271°C, allowing the reaction to continue stably. The gas after condensate separation remained stable at 95% methane, 4% hydrogen, and 1% carbon dioxide, yielding a high-purity methane gas suitable for use as a raw material for city gas. The ethanol conversion rate remained at 100%.

[Comparative Example 5]
The test was conducted in the same manner as in Example 7, except that the gas supplied to the catalyst layer was a mixture of ethanol (200 mL/min), hydrogen (200 mL/min), and water vapor (400 mL/min).

Table 15 shows the steam reforming catalyst inlet temperature, the upper, middle, and lower temperatures of the steam reforming catalyst layer, the upper, middle, and lower temperatures of the methane catalyst layer, and the analysis results of the gas composition after condensate separation, as well as the ethanol conversion rate, at 1, 2, 3, and 4 hours after the start of the reaction.

The ethanol conversion rate remained stable at 100%. The gas composition after condensate separation remained stable at 85.5% methane, 1.9% hydrogen, and 12.6% carbon dioxide. Although the obtained gas is primarily methane, its high carbon dioxide concentration makes it difficult to use as a raw material for city gas without some kind of carbon dioxide removal process.

[Other Embodiments]
Finally, other embodiments of the method for producing high-calorific value fuel gas according to the present invention will be described. Note that the configurations disclosed in each of the following embodiments can be applied in combination with configurations disclosed in other embodiments, as long as no inconsistencies arise.

In the embodiments of Examples 1 to 6, examples in which the steam reforming reactor used in the steam reforming process is configured as an adiabatic reactor were particularly described. However, the reactor used in the steam reforming process according to the present invention is not particularly limited and may be configured as a heat exchange reactor.

In the embodiments of Examples 1 to 6, examples were particularly described in which the steam reforming reactor and the methane reactor used in the methane process are housed in separate reactors, and further equipped with a heat exchanger, so that the gas exiting the steam reforming reactor is cooled by passing it through the heat exchanger before being inserted into the methane reactor. However, the steam reforming reactor and the methane reactor according to the present invention may be configured as a single unit. For example, by filling a single reactor with the same or different catalysts, and by not providing a heat exchanger near the inlet, or by designing the heat exchange capacity near the inlet to be smaller than that near the outlet, the reaction may proceed under adiabatic or near-adiabatic conditions in the first half of the reactor, and under isothermal or near-adiabatic conditions in the second half. In this case, the raw material gas is introduced into the steam reforming reactor at 300°C. The reaction proceeds adiabatically in the first half, with the catalyst layer temperature peaking at, for example, around 400°C. In the second half, the gas produced by the steam reforming reaction is gradually cooled, and the methane reaction proceeds, with the reactor exiting at a temperature between 230°C and 330°C. In this scenario, the steam reforming process takes place in the first half of the reactor, and the methane reaction process is carried out in the second half.

Furthermore, the present invention does not exclude the use of conventionally known methods for improving methane conversion rates. For example, a configuration may be used in which the reaction gas is cooled in an intermediate stage of a multi-stage reaction to condense and separate a portion of the water vapor. Even in this case, the method of the present invention can reduce the concentrations of hydrogen and carbon dioxide in the fuel gas without excessively removing water vapor.

Furthermore, the configurations disclosed in the above embodiments (including other embodiments, the same applies hereinafter) can be applied in combination with configurations disclosed in other embodiments, provided that no inconsistencies arise. Moreover, the embodiments disclosed herein are illustrative, and the embodiments of the present invention are not limited thereto; they can be modified as appropriate without departing from the purpose of the present invention.

This invention can be used, for example, as a method for producing fuel gas for supply as city gas.

1: Raw material gas adjustment section 2: Steam reforming reaction section 3: Methanation reaction section 4a: Selective oxidation reaction section 4b: Dehydration and hydrogenation reaction section

Claims (5)

A process of preparing a raw material gas by adding hydrogen and water vapor to ethanol such that the molar ratio of hydrogen/ethanol is 2.0 or more and 2.5 or less, and the molar ratio of water vapor/ethanol is 1.2 or more and 4 or less.
A steam reforming step in which the aforementioned raw material gas is brought into contact with a catalyst containing ruthenium or nickel at a temperature of 300°C to 700°C,
The process includes a methanation step in which the gas obtained in the steam reforming step and cooled is brought into contact with a catalyst containing ruthenium or nickel under conditions that the catalyst outlet temperature is between 230°C and 330°C.
A method for producing a high-calorific value fuel gas , wherein the catalyst outlet temperature of the steam reforming step is higher than the catalyst outlet temperature of the methane step . The method for producing a high-calorific value fuel gas according to claim 1, wherein the steam reforming step is carried out under substantially adiabatic conditions with a catalyst inlet temperature of 300°C or higher, and the gas obtained in the steam reforming step is cooled to 230°C or higher and then to 330°C or lower before being supplied to a heat exchange reactor to carry out a methane step. A method for producing a high-calorific value fuel gas according to claim 1 or 2, further comprising: adding an amount of oxygen to a fuel gas obtained in a methanation step, which is mainly composed of methane and contains hydrogen, such that the molar ratio of oxygen to hydrogen in the fuel gas is 0.24 to 0.45; and then reacting the oxygen with the hydrogen in the fuel gas in the presence of a selective oxidation catalyst capable of selectively oxidizing hydrogen. A method for producing a high-calorific value fuel gas according to claim 1 or 2, further comprising: adding an amount of ethanol to a fuel gas obtained in a methanation step, such that the molar ratio to hydrogen in the fuel gas is 0.45 or more and 0.9 or less; and then reacting the ethanol with the hydrogen in the fuel gas in the presence of a dehydration hydrogenation catalyst to obtain ethane. The method for producing a high-calorific value fuel gas according to claim 1 or 2, wherein the high-calorific value fuel gas contains methane or methane and ethane, and the proportion of methane or methane and ethane is 93% or more on a volume basis after dehydration.