WO2018124791A1 - Heat exchange microchannel reactor for oxidative coupling of methane - Google Patents

Abstract

The present invention can provide a method, which: provides a means capable of effectively controlling reaction heat by using a heat exchange microchannel reactor so as to simultaneously carry out an exothermic reaction, of oxidative coupling of methane, and an endothermic reaction, thereby giving and receiving reaction heat, and thus can obtain a high C2 product selectivity; and improves thermal efficiency of a process by achieving thermally neutral reaction conditions, and thus can more efficiently convert methane.

Description

Heat Exchange Microchannel Reactor for Oxidative Dimerization of Methane

According to the present invention, the methane oxidative dimerization reaction channel and the endothermic reaction flow path of the exothermic reaction are alternately provided, and the temperature (T ₁ ) of the exothermic reaction flow path is higher than the temperature (T ₂ ) of the endothermic reaction flow path from the exothermic reaction path to the endothermic reaction flow path. Heat exchange microchannel reactor for heat transfer; And to converting methane in the heat exchange microchannel reactor to produce a gas product.

The present invention also relates to a method for more efficiently converting methane by combining an oxidative dimerization reaction of methane which is an exothermic reaction and a steam reforming reaction of methane which is an endothermic reaction.

Methane, the main component of natural gas, is mainly used for power generation or domestic heat sources, but some are used as raw materials for producing liquid fuels or chemicals. Methane conversion reaction for this is mostly through the reforming reaction of methane, a synthesis gas manufacturing process. These syngases are used as raw materials for methanol synthesis, hydrogen production, ammonia production or synthetic oil production by Fischer-Tropsch reaction.

On the other hand, the oxidative coupling of methane (methane) has a merit as a reaction that is directly converted to ethylene and ethane by the reaction of methane and oxygen on the catalyst, but accompanied by various side reactions. Methane and oxygen react to form ethylene or ethane according to the forward reaction as in Scheme 1) or Scheme 2). In this process, very strong exothermic reaction may occur with side reactions such as Scheme 3) and Scheme 4). . If the reaction temperature is not easy to control the reaction heat of the reaction forward, the reaction proceeds further with CO and CO ₂ to reduce the yield of the desired C ₂ hydrocarbon compound. Therefore, control of the heat of reaction in the oxidative dimerization of methane is important for the stability of the reaction process as well as the yield of the product.

CH ₄ + ¼ O _2- > ½ C ₂ H ₆ + ½ H ₂ O, ΔH ₂₉₈ = -87.8 kJ / mol 1)

CH ₄ + ½ O _2- > ½ CH ₂ = CH ₂ + H ₂ O, ΔH ₂₉₈ = -140.4 kJ / mol 2)

CH ₄ + 3/2 O _2- > CO + H ₂ O, ΔH ₂₉₈ = -518.7 kJ / mol 3)

CH ₄ + 2 O _2- > CO ₂ + 2 H ₂ O, ΔH ₂₉₈ = -801.3 kJ / mol 4)

On the other hand, methane reforming reaction is a strong endothermic reaction, as shown in the following reaction scheme, unlike methane oxidative dimerization reaction. That is, a lot of heat must be supplied from the outside in order for methane reforming to occur.

CH ₄ + H ₂ O-> CO + 3H ₂ ΔH ₂₉₈ = 206.2 kJ / mol 5)

CH ₄ + CO _2- > 2CO + 2H ₂ ΔH ₂₉₈ = 247.2 kJ / mol 6)

3CH ₄ + 2H ₂ O + CO _2- > 4CO + 8H ₂ ΔH ₂₉₈ = 659.6 kJ / mol 7)

CH ₃ CH _3- > CH ₂ = CH ₂ + H ₂ ΔH ₂₉₈ = 173 kJ / mol 8)

In the reforming reaction of methane, methane may react with water vapor as in Scheme 5), carbon dioxide may be used instead of water vapor as in Scheme 6), and steam and carbon dioxide may be used at the same time as in Scheme 7).

As in Scheme 8, the cracking reaction of ethane is also endothermic.

Methane is a very stable compound, and when it is converted to another compound, there is a problem that thermal efficiency is lowered because it involves a very strong endothermic reaction or exothermic reaction. Until now, research on each of the above reactions has been conducted a lot, but there are very few examples of implementing both reactions in the same reactor.

In general, a microchannel reactor has a plate stacked structure, in which a catalyst layer has a cooling layer above and below, and a catalyst layer has a thickness of several millimeters. The microchannel reactor has attracted attention as a strong exothermic or endothermic reactor because of its excellent heat exchange performance since the heat transfer area is 10 to 100 times larger than the conventional tube-type fixed bed catalytic reactor. Microchannel reactors are currently being commercialized as compact reactors of the Fischer-Tropsch reaction, a highly exothermic reaction.

The inventors of the present invention implement reactions in which the entire reaction is thermally neutralized using a reactor having excellent heat exchange in one reactor, and the products of these reactions are recycled to each reaction of the unreacted product through a secondary reaction. It was confirmed that the efficiency of the and completed the present invention. Therefore, the present invention is to improve the thermal efficiency of the reaction by performing the reaction in the thermal neutral conditions when the conversion from methane to other compounds, and to provide a high efficiency methane conversion method through integration with the secondary reaction.

The present invention also provides a microchannel reactor in which a strong exothermic reaction layer and a strong endothermic reaction catalyst layer are separated from each other and alternately stacked, and easily control heat of reaction and have high thermal efficiency in a reaction process.

The first aspect of the present invention is provided with an exothermic reaction flow path and an endothermic reaction flow path, wherein the temperature (T ₁ ) of the exothermic reaction flow path where the oxidative dimerization reaction of methane is performed is performed in the endothermic reaction flow path where the steam reforming reaction of methane is performed. In a heat exchange reactor in which heat transfer from the exothermic reaction passage to the endothermic reaction passage is higher than the temperature T ₂ , the first step of performing oxidative dimerization of methane in the exothermic reaction passage and performing steam reforming reaction of methane in the endothermic reaction passage It provides a method of converting methane characterized in that it comprises a.

In this case, ethylene and / or ethane are formed as a product as C2 hydrocarbons by the oxidative dimerization reaction of methane in a first step, and a synthesis gas is formed as a product by steam reforming of methane, optionally, formed in the first step. A second step of preparing synthetic oil from the synthesis gas by methanol synthesis, hydrogen production, ammonia production or a Fischer-Tropsch reaction; A third step of preparing a liquid hydrocarbon product by ethylene oligomerization reaction from the C2 hydrocarbon formed in the first step; And a fourth step of recycling the unreacted gas of the third step to the steam reforming reaction of the first step and the oxidative dimerization reaction of methane.

The second aspect of the present invention includes one or more exothermic reaction passages and two or more endothermic reaction passages, wherein the temperature (T ₁ ) of the exothermic reaction passage is higher than the temperature (T ₂ ) of the endothermic reaction passage and endothermic from the exothermic reaction passage. In a heat exchange microchannel reactor in which heat is transferred to a reaction flow path, an exothermic reaction flow path is filled with a catalyst for oxidative dimerization reaction of methane, an endothermic reaction flow path is filled with a catalyst for endothermic reaction, and exchanges heat with an endothermic reaction flow path. Heat exchange microchannel reactor characterized in that the reaction temperature in the exothermic reaction passage is controlled within the range of 800 ℃ ± 50 ℃, the reaction temperature in the endothermic reaction passage through the heat exchange with the exothermic reaction passage is controlled within the range of 750 ℃ ± 50 ℃ To provide.

The third aspect of the present invention includes one or more exothermic reaction passages and two or more endothermic reaction passages, wherein the temperature T ₁ of the exothermic reaction passage is higher than the temperature T ₂ of the endothermic reaction passage and endothermic from the exothermic reaction passage. In a heat exchange microchannel reactor in which heat is transferred to a reaction passage, an exothermic reaction passage is filled with a catalyst for oxidative dimerization reaction of methane, an endothermic reaction passage is filled with a catalyst for endothermic reaction, and the calorific value of the exothermic reaction passage can be controlled. In order to provide a heat exchange microchannel reactor, the thickness of a catalyst bed in an exothermic reaction flow path located between two adjacent endothermic flow paths is controlled within a range of 1 to 5 mm.

In this case, the thickness of the catalyst bed in the endothermic reaction flow path may be adjusted within 0.1 to 2 times the thickness of the catalyst layer in the exothermic reaction flow path so as to remove the calorific value downstream in the exothermic reaction flow path.

The fourth aspect of the present invention includes one or more exothermic reaction passages and two or more endothermic reaction passages, and the temperature (T ₁ ) of the exothermic reaction passage is higher than the temperature (T ₂ ) of the endothermic reaction passage and endothermic from the exothermic reaction passage. In a heat exchange microchannel reactor in which heat is transferred to a reaction flow path, an exothermic reaction flow path is filled with a catalyst for oxidative dimerization reaction of methane, an endothermic reaction flow path is filled with a catalyst for reforming reaction of methane, and The reforming reaction provides a heat exchange microchannel reactor characterized by controlling the conversion of methane to 95% or less of equilibrium conversion under reaction conditions by controlling catalytic activity.

The fifth aspect of the present invention includes one or more exothermic reaction passages and two or more endothermic reaction passages, wherein the temperature T ₁ of the exothermic reaction passage is higher than the temperature T ₂ of the endothermic reaction passage and endothermic from the exothermic reaction passage. In a heat exchange microchannel reactor in which heat is transferred to a reaction flow path, an exothermic reaction flow path is filled with a catalyst for oxidative dimerization reaction of methane, an endothermic reaction flow path is filled with a catalyst for reforming reaction of methane, and upstream of the exothermic reaction flow path. The methane conversion rate is 60% to 95% in the reforming reaction of methane in the endothermic passage to lower the rate of oxidative dimerization of the methane to suppress rapid temperature increase and to remove the calorific value of the oxidative dimerization reaction of methane downstream in the exothermic passage. It provides a heat exchange microchannel reactor characterized by being controlled in a range.

The sixth aspect of the present invention provides a method for producing a gas product by converting methane or ethane in the heat exchange microchannel reactor of the second to fifth aspects.

In the heat exchange microchannel reactors of the second to fifth embodiments, the catalyst packed in the endothermic flow path may be a catalyst for reforming the methane, and the synthesis gas may be prepared by reforming the methane in the endothermic flow path. In addition, in the heat exchange microchannel reactor of the second to fifth embodiments, the catalyst packed in the endothermic flow path may be a catalyst for dehydrogenation of ethane, and ethylene may be prepared through pyrolysis of ethane without additional catalyst filling. have.

The present invention achieves the thermal neutral reaction conditions by performing heat exchange reaction of methane which is exothermic reaction and steam reforming reaction of methane which is endothermic at the same time to improve the thermal efficiency of the process, and through integration with the rear end reaction It can provide a high efficiency whole integrated process.

In addition, the present invention provides a means for effectively controlling the heat of reaction by exchanging and receiving the heat of reaction by simultaneously carrying out the oxidative dimerization reaction and endothermic reaction of methane exothermic reaction using a heat exchange microchannel reactor, thereby selecting a high C2 product It is possible to provide a way to more efficiently convert methane by improving the thermal efficiency of the process by obtaining a diagram and achieving thermal neutral reaction conditions.

1 is a schematic of a process for preparing a liquid hydrocarbon or compound in methane.

2 is a process schematic diagram illustrating the conversion process of the methane-containing gas in more detail.

FIG. 3 is a schematic diagram (a) of a microchannel reactor in which methane oxidative dimerization (OCM) and steam reforming (SMR) of methane occurs, in-channel exothermic and endothermic reactions, and heat exchange between channels; The three-dimensional drawing (b) (c) shows a photograph (d) of the microchannel reactor actually manufactured.

Figure 4 shows the parallel conversion rate according to the reaction temperature of the methane reforming reaction under various reaction conditions

5 (a) to (e) are photographs of the microchannel reactors produced by the respective production examples.

The methane conversion method of the present invention is characterized in that the thermal neutral reaction is performed by exchanging heat of reaction by performing oxidative dimerization reaction of methane and steam reforming reaction of methane in one reactor.

To this end, the methane conversion method of the present invention comprises the exothermic reaction flow path and the endothermic reaction flow path of the oxidative dimerization reaction of methane and the steam reforming reaction of methane, and the temperature of the exothermic reaction flow path where the oxidative dimerization reaction of methane is performed (T ₁ ) is higher than the temperature (T ₂ ) of the endothermic reaction flow path where the steam reforming reaction of methane is being performed, characterized in that it is carried out in a heat exchange reactor in which heat transfer from the exothermic reaction flow path to the endothermic reaction flow path. In a preferred embodiment of the present invention, the heat exchange reactor may be a microchannel reactor, and each channel of the microchannel reactor may be filled with a catalyst for oxidative dimerization of methane or a catalyst for steam reforming of methane.

The methane conversion method according to the first aspect of the present invention is characterized in that in the heat exchange reactor, a oxidative dimerization reaction of methane is carried out in an exothermic passage and a steam reforming reaction of methane is carried out in an endothermic reaction passage. It is done. At this time, each reaction occurs while passing through the catalyst layer in the channel of the heat exchange reactor.

The reactant injected into the heat exchange reactor of the first stage does not need to be 100% methane, which contains methane as a main component, and may be derived from natural gas or petrochemical by-products, and furthermore, a small amount of ethane, propane or nitrogen or carbon dioxide. It may include.

Another steam reforming reaction of the methane of the first step is a reaction of reacting methane with steam, carbon dioxide, or a mixture thereof to generate syngas (CO + H ₂ ). Steam reforming of methane occurs at 700–900 ° C., a reaction temperature similar to that of methane oxidative dimerization, but a strong endothermic reaction occurs.

The steam reforming reaction of methane is preferably carried out at a reaction pressure of 1 to 30 bar and a molar ratio of water vapor / methane of 1 to 4. Since the ratio of H ₂ / CO of syngas obtained in steam reforming of methane is higher than 3, in the present invention, subsequent reactions utilizing these syngases are either methanol synthesis or Fischer-Tropsch reactions. In order to lower the H ₂ / CO ratio H ₂ O / CH ₄ ratio was lowered and carbon dioxide was added to some reactions. For example, _{CH 4: H 2 O: CO} 2 molar ratio is 1: 1.5 ~ 2.0: 0.4 ~ 0.8 when a is the reaction temperature is 800 ~ 850 ℃, the H ₂ / CO molar ratio of 2.0 ~ 2.5 degree methanol synthesis or Syngas suitable for Fischer-Tropsch reactions can be obtained.

In the present invention, the oxidative dimerization reaction of methane and the steam reforming reaction of methane are important for the transfer of reaction heat in close contact with each other in the same reactor, and it is preferable to use a microchannel reactor for this purpose. The microchannel reactor may be suitable for a reaction having a large heat transfer area relative to the volume of the reactor.

Meanwhile, the heat exchanging microchannel reactor according to the present invention includes one or more exothermic reaction passages and two or more endothermic reaction passages, and the exothermic reaction passage temperature T ₁ is higher than the temperature of the endothermic reaction passage T ₂ . Heat is transferred from the reaction passage to the endothermic reaction passage, the exothermic reaction passage is filled with a catalyst for oxidative dimerization reaction of methane, and the endothermic reaction passage is filled with a catalyst for endothermic reaction. The exothermic reaction flow path and / or the endothermic reaction flow path may be charged with a particulate catalyst to form a catalyst layer. However, in the endothermic reaction flow path, a catalyst-free reaction (eg, a catalyst cracking / pyrolysis reaction of ethane) may be performed without additional catalyst filling, which is also within the scope of the present invention.

In the heat exchanging microchannel which simultaneously performs the exothermic and endothermic reactions according to the present invention, the endothermic reaction can absorb the heat of the exothermic reaction so that the endothermic reaction can be carried out. Preferably within ± 50 ° C., more preferably within ± 50 ° C., even more preferably within ± 30 ° C.).

The endothermic reaction is not particularly limited, but when the exothermic reaction is oxidative dimerization of methane, the endothermic reaction may be a reforming reaction of methane and / or a dehydrogenation or pyrolysis reaction of ethane.

In the present invention, it was possible to effectively control the strong exotherm of the oxidative dimerization reaction of methane by using a heat exchange microchannel reactor capable of heat transfer to the endothermic reaction flow path, specifically, a catalyst for oxidative dimerization reaction of methane in one layer of the microchannel reactor The exothermic reaction was carried out to fill, and it was confirmed that endothermic reactions such as reforming reactions of methane having catalytic activity controlled therebetween were effectively controlled to control the heat of reaction of the oxidative dimerization reaction of methane.

The optimum reaction temperature for the oxidative dimerization reaction of methane is around 800 ℃, which is somewhat different depending on the catalyst, but the yield of C ₂ hydrocarbon compound is maximized near 800 ℃, and below 750 ℃, most of the methane oxidative dimerization catalyst There is little activity, and the selectivity of the C ₂ product is greatly reduced above 900 ° C.

Methane reforming, an example of endothermic reaction, is a reaction in which methane is reacted with steam, carbon dioxide, or a mixture thereof to produce syngas (CO + H ₂ ). The reforming reaction of methane occurs at 600 ~ 900 ℃, which is similar to the oxidative dimerization reaction of methane, but a strong endothermic reaction occurs. The methane reforming reaction can achieve high methane conversion around 700-800 degrees. The reforming reaction of methane is a reversible reaction and the equilibrium conversion increases with increasing reaction temperature. As shown in FIG. 2, when the molar ratio of H ₂ O / CH ₄ is 3, the reaction pressure is 1 atm, and the reaction temperature is 800 ° C., the conversion rate of methane is 99.8%, and when the ratio is 1 at the same pressure, the methane conversion rate is 90.1. At the rate of 5 atm, the conversion rate drops significantly to 68.3%. At 800 ° C., which is the optimum reaction temperature of the oxidative dimerization of methane, the difference in the methane conversion rate of the methane reforming reaction is too large as described above, which may cause a problem in controlling the exotherm of the methane oxidative dimerization reaction due to the endotherm of the methane reforming reaction. That is, general reaction temperature of methane reforming reaction is around 700 ℃, while the reaction temperature of oxidative dimerization reaction of methane is 800 ℃, it is difficult to control the exotherm of the oxidative dimerization reaction of methane.

Therefore, in order to solve this problem, according to the present invention, the heat exchange microchannel reactor capable of heat transfer to the endothermic reaction passage is controlled in the exothermic reaction passage through heat exchange with the endothermic reaction passage in the reaction temperature within the range of 800 ° C. ± 50 ° C. The reaction temperature in the endothermic reaction passage through heat exchange with the reaction passage is characterized in that it is controlled within the range of 750 ℃ ± 50 ℃.

In the present invention, the oxidation dimerization reaction of methane, the reaction temperature is 700 ~ 900 ℃, preferably 800 ℃ ± 50 ℃, the reaction pressure 1 ~ 10 bar, methane / oxygen molar ratio is 2 ~ 10, space velocity is 1000 ~ 50000 h ^-1 .

Temperature control of the oxidative dimerization reaction layer of methane in the microchannel reactor according to the present invention may be achieved by controlling the injection amount of the reactant of the steam reforming reaction of methane compared to the oxidative dimerization reactant. As the reactants of the steam reforming reaction of methane increases, the endothermic reaction may increase and the temperature of the oxidative dimerization reaction of methane may drop. Combining the oxidative dimerization reaction of methane and the steam reforming reaction of methane as in the present invention has the advantage that the waste heat of the oxidative dimerization reaction of methane can be utilized as the reaction heat of the methane reforming reaction, the reaction heat of the oxidative dimerization reaction of methane There is an easier advantage to solve the problem of the prior art, which requires the use of much more high temperature steam in order to control the exothermic amount of the high temperature reaction.

However, when increasing the reactants of the methane reforming reaction to lower the temperature of the methane oxidative dimerization reaction bed, the methane conversion of the methane reforming reaction is too high at the top of the catalyst and less unreacted methane at the bottom. Therefore, the problem of control of reaction heat becomes difficult at the lower end of the catalyst of the oxidative dimerization reaction.

In order to solve the above problems, according to one embodiment of the present invention, the plates (plates) are stacked vertically at regular intervals, each reaction is alternately made in the channel consisting of plates Each channel used a microchannel reactor filled with a catalyst (see FIG. 3 (b)). The catalytic reaction layer of each channel may be an exothermic reaction flow path or an endothermic reaction flow path. At this time, the oxidative dimerization reaction channel of methane which is exothermic reaction and the methane reforming reaction channel which is endothermic reaction are laminated | stacked alternately by layer, and the methane reforming reaction layer of methane exists in the lower part and the upper part of methane oxidization dimerization reaction layer, Heat transfer can be achieved by sending and receiving reaction heat up and down the reaction layer (see FIG. 3 (a)).

The present invention was adjusted so that the reaction conditions of the endothermic reactions performed adjacent to the reaction heat of the exothermic reaction correspond to the conditions of the exothermic reaction. For this purpose, it has been found that the reactor structure, the catalytic performance control, and the reaction condition control in the endothermic reaction are necessary for the exothermic reaction heat control.

That is, through the experiments the present invention proposes a reaction heat control method of the exothermic reaction in the heat exchange microchannel reactor.

The first method relates to the structure of a microchannel reactor, in which the contact time of the reactants of the methane reforming reaction layer is controlled by controlling the thickness of the methane reforming catalyst layer relative to the thickness of the catalytic reaction layer of the oxidative dimerization reaction of methane. It is a way. The thickness of the oxidation dimerization catalyst layer of methane is suitably 1 to 5 mm. When the thickness of the catalyst layer is too thin, the amount of catalyst charged in the reactor is small, so economic efficiency is low, and when the thickness is too thick, it is difficult to control the amount of heat generated, so the above range is preferable. The thickness of the reforming catalyst layer of methane depends on the thickness of the oxidative dimerization catalyst layer of methane, and preferably 0.1 to 2 relative to the thickness of the oxidative dimerization catalyst layer of methane. If the thickness of the methane reforming catalyst layer is less than the ratio of 0.1, the contact time of the reactants is too short, the conversion rate of methane in the methane reforming reaction is too low, it may be difficult to effectively remove the heat of the oxidative dimerization reaction of methane, the ratio If it is larger than 2, the contact time is increased in the methane reforming reaction, so that the conversion of the methane reforming reaction may occur only at the upper portion of the catalyst, and thus, it may be difficult to control the heat of reaction of the oxidative dimerization reaction of the middle and lower methane. Therefore, the catalyst layer thickness ratio is suitable for controlling the heat of reaction of the oxidative dimerization reaction of methane.

The second method of controlling the heat of reaction of the oxidative dimerization reaction of methane is a method of artificially reducing the activity of the catalyst in the catalyst of the methane reforming reaction. The reforming catalyst of methane may be one in which catalytically active metal components such as Ni, Pt, Rh, Co and the like are supported on various catalyst supports. Methods of lowering the artificial catalyst activity include a catalyst sintering method at a high temperature (> 1000 ° C.), supporting an alkali metal such as K, and catalyst poisoning by a component such as S, and the like, but the present invention is not particularly limited thereto. . The catalytic activity of the methane reforming reaction is lowered to achieve an optimum temperature similar to that of the methane reforming reaction. When the methane conversion rate of methane reforming is lowered to 95% or less, as mentioned above, rapid methane conversion at the top of the catalyst layer of methane reforming can be prevented, thereby preventing the temperature of the methane oxidative dimerization reaction layer from increasing rapidly. .

The third method is to control the conversion of methane by adjusting the reaction conditions of the methane reforming reaction. The reforming reaction of methane can control the rate of methane conversion with temperature by the ratio of water vapor or carbon dioxide or a mixture thereof or the reaction pressure. As shown in FIG. 2, the conversion rate of methane is lower as the molar ratio of H ₂ O / CH ₄ or CO ₂ / CH ₄ or (H ₂ O + CO ₂ ) / CH ₄ decreases, and as the reaction pressure increases. In the present invention, the reaction conditions were adjusted to prevent excessive conversion in the upper stage of the catalyst layer of the methane reforming reaction. The proper methane conversion of methane reforming is thus 60% to 95%. If the conversion rate is lower than 60% may cause a problem that the efficiency of the after-stage reaction of the produced synthesis gas is lowered, and if the conversion rate is higher than 95%, the temperature in the bottom catalyst layer of the oxidative dimerization reaction of methane may be too high. This occurs because almost no endothermic reaction occurs at the bottom of the methane reforming reaction.

The reaction heat control method of the oxidative dimerization reaction of methane can be used without limitation to the type of endothermic reaction. The dehydrogenation or pyrolysis reaction of ethane is also endothermic and is preferred as endothermic because the reaction can be carried out at temperature conditions similar to the oxidative dimerization of methane. Thus, non-limiting examples of endothermic reactions include methane reforming as well as ethane cracking.

Therefore, according to the present invention, an exothermic reaction passage filled with a catalyst for oxidative dimerization reaction of methane and an endothermic reaction passage filled with an endothermic catalyst are alternately provided, and the temperature (T ₁ ) of the exothermic reaction passage is endothermic. The heat exchange microchannel reactor, which is heat transfer from the exothermic reaction flow path to the endothermic reaction flow path higher than the temperature T _{2 of the} flow path, is characterized by satisfying each of the following four requirements or a combination thereof:

(i) The reaction temperature in the exothermic reaction passage through heat exchange with the endothermic reaction passage is controlled within the range of 800 ° C. ± 50 ° C., and the reaction temperature in the endothermic reaction passage through the heat exchange with the exothermic reaction passage ranges from 750 ° C. ± 50 ° C. Controlling within;

(ii) the thickness of the catalytic bed in the exothermic reaction flow path located between two adjacent endothermic reaction flow paths (ie, the two endothermic reaction flow paths) so as to enable control of the calorific value in the exothermic reaction flow path without lowering the product yield. The thickness of the catalytic bed in the endothermic reaction flow path is adjusted to within the range of 1 to 5 mm and / or to remove the calorific value downstream in the exothermic flow path. Adjusting within a range of 0.1 to 2 times the thickness of the catalyst layer in the flow path;

(iii) reforming the methane in the endothermic reaction flow path using an endothermic catalyst having a catalytic activity reduced to 95% or less of the methane conversion to equilibrium conversion under the reaction conditions by controlling the catalytic activity; And

(iv) Reforming the methane in the endothermic passage to slow down the rate of oxidative dimerization of the methane upstream in the exothermic flow passage to inhibit rapid temperature increase and to remove the calorific value of the oxidative dimerization reaction of methane downstream in the exothermic flow passage. To control the methane conversion in the range of 60% to 95%.

In the heat exchange microchannel reactor according to the present invention, the fluid flow in the exothermic reaction flow path and the fluid flow in the endothermic reaction flow path are preferably in the same direction. For example, if the fluid flow of the oxidative dimerization reaction of methane and the flow of the reforming reaction of methane are reversed, the product yield is lowered due to the high temperature in the upper catalyst layer of the oxidative dimerization reaction of methane (Comparative Example 6).

The heat exchange microchannel reactor according to the present invention can be used in a methane conversion process. For example, the methane conversion method of the present invention may perform the oxidative dimerization of methane in the exothermic flow passage in the heat exchange microchannel reactor, and / or the reforming reaction of methane in the endothermic flow passage. At this time, each reaction occurs while passing through the catalyst layer in the channel of the heat exchange reactor.

Accordingly, one aspect of the present invention provides a method for producing a gas product by converting methane in the heat exchange microchannel reactor of the various aspects of the present invention.

In this case, by performing the oxidative dimerization reaction of methane in the exothermic flow path can be converted to a hydrocarbon of C2 or more including ethylene and / or ethane from a methane-containing gas to produce a hydrocarbon of C2 or more. In addition, synthesis gas may be prepared from methane-containing gas by reforming the methane in the endothermic reaction flow path, or ethylene may be produced from the ethane by carrying out dehydrogenation of ethane or cracking of ethane in the endothermic flow path. In the exothermic reaction passage in the heat exchange microchannel reactor according to the present invention, the reaction temperature is controlled within the range of 800 ℃ ± 50 ℃, the reaction temperature in the endothermic reaction passage is controlled within the range of 750 ℃ ± 50 ℃, methane in the exothermic reaction passage The ethane-containing gas formed through the oxidative dimerization reaction may be introduced as a reactant into the endothermic reaction path that performs dehydrogenation of ethane or cracking reaction of ethane, and at this time, the preheating of the reactants may be omitted during the endothermic reaction. Can be.

The products of the oxidative dimerization of methane may include ethane, ethylene, CO, CO ₂ , H ₂ , H ₂ O, traces of C 3+ hydrocarbons and unreacted methane. Preferably ethylene and / or ethane can be formed as the product as C ₂ hydrocarbons by oxidative dimerization of methane. The production rate of ethane and ethylene is similar, but as the reaction temperature increases, the production rate of ethylene tends to increase. The oxidative dimerization of methane is a strong exothermic reaction. If the heat of reaction is not well controlled, the production of CO or CO ₂ increases and the yield of C2 + hydrocarbons decreases. Therefore, it is important to keep the reaction temperature within a certain temperature range. To this end, in the present invention, the temperature (T ₁ ) of the exothermic reaction passage in which the oxidative dimerization reaction of methane is performed while the exothermic reaction passage and the endothermic reaction passage are alternately arranged adjacent to each other is the temperature of the endothermic reaction passage (T ₂ ). Not only is it performed in a heat exchanging microchannel reactor that is heat transfer from the exothermic reaction passage to the endothermic reaction passage, the heat exchanging microchannel reactor may further meet the characteristics of various combinations of (i) to (iv).

In addition, the heat exchange microchannel reactor according to the present invention can perform the methane conversion process through the reforming reaction of methane in the endothermic reaction passage. The reactants of the methane reforming reaction are methane; And water vapor, carbon dioxide or mixtures thereof. In addition, a portion of the product containing the unreacted methane of the oxidative dimerization of methane may be recycled to the reactant in the endothermic flow path for performing the methane reforming reaction. By-products of methane oxidative dimerization include CO, H ₂ , H ₂ O and CO ₂ , which are products or reactants of methane reforming, and it is advantageous to recycle them to methane reforming rather than oxidative dimerization. Do. In conclusion, the amount and composition of the by-products in the oxidative dimerization of methane depends on the reaction conditions, but it is advantageous in terms of the efficiency of the overall process to recycle the remaining by-products to the oxidative dimerization reaction of methane.

In the present invention, the oxidative dimerization reaction of methane and the reforming reaction of methane are important for the transfer of reaction heat in close contact with each other in the same reactor. To this end, in the present invention, a microchannel reactor is used as the heat exchange reactor, and each channel of the microchannel reactor may be filled with a catalyst for oxidative dimerization of methane or a catalyst for reforming of methane. Methane conversion rate in methane reforming in the endothermic flow path to lower the rate of oxidative dimerization of methane upstream in the exothermic flow path to suppress rapid temperature increase and to remove the calorific value of the oxidative dimerization reaction of methane downstream in the exothermic flow path. Can be adjusted within the range of 60% to 95%. The reforming reaction of methane may control the conversion of methane by controlling the molar ratio of water vapor / methane, the molar ratio of carbon dioxide / methane or the molar ratio of (water vapor + carbon dioxide) / methane in the endothermic reaction channel. The reforming reaction of methane is preferably adjusted to a molar ratio of (water vapor + carbon dioxide) / methane to 0.8 to 2 in order to control the methane conversion to 60 to 95%. Therefore, the reforming reaction of methane is preferably carried out at a reaction pressure of 1 ~ 10bar, a molar ratio of (water vapor + carbon dioxide) / methane of 0.8 ~ 2, the space velocity of 1000 ~ 50000 h ^-1 . In addition, by controlling the catalytic activity of the methane reforming reaction, the conversion of methane to the equilibrium conversion in the reaction conditions can be adjusted to 95% or less.

Subsequent reactions utilizing syngas obtained for the reforming of methane may be methanol synthesis or Fischer-Tropsch reactions. The reactant composition of the methane reforming reaction can be adjusted to match the composition of the syngas suitable for each subsequent reaction.

Since the product of the oxidative dimerization reaction of methane prepared in the heat exchange microchannel reactor according to the present invention is a high temperature of 700 ° C or more, it has to be rapidly cooled below a certain level. This is because a highly reactive ethylene product can undergo a secondary reaction if it is maintained at a high temperature for a long time. Cooling of the hot product can be accomplished by cooling water or steam. The high pressure or high temperature steam generated in this process can be utilized as utility steam in the process. The cooled reaction product is further secondary cooled and then fed to a gas-liquid separator to separate the gas-liquid separation to recover the liquid product, water, and the gas product is fed to the olefin oligomerization reactor, which is a subsequent reactor after being preheated. Can be.

Alternatively, all of the products of the oxidative dimerization reaction of methane may be introduced into a subsequent olefin oligomerization reactor without performing the secondary cooling and gas-liquid separation separately. At this time, the temperature of the product flow after the primary cooling may be adjusted in the range of 250 ℃ to 500 ℃ temperature of the ethylene oligomerization reaction. When the reaction product of the oxidative dimerization reaction of methane is injected directly into the ethylene oligomerization reactor without secondary cooling, the process is simplified by eliminating the secondary cooler, the gas-liquid separator, and the preheater, and thus, no separate preheating is necessary. While there is an advantage in improving the thermal efficiency of the process, there may be an effect on water in the catalyst of the later ethylene conversion reactor.

According to a first aspect of the present invention, after the first step, a second step of preparing a synthetic oil by methanol synthesis, hydrogen production, ammonia production, or a Fischer-Tropsch reaction is carried out from the synthesis gas formed in the first step. It may further include.

The first aspect of the present invention may further include a third step of preparing a liquid hydrocarbon product by ethylene oligomerization reaction from the C ₂ hydrocarbon formed in the first step after the second step. The composition discharged through the reactor of the third stage may include CO, CO ₂ , H ₂ O, hydrogen, methane, C ₂ hydrocarbons, C ₃ ~ C ₄ hydrocarbons, C 5 + hydrocarbons, aromatics (aromatic) and the like.

The C ₂ hydrocarbons discharged through the reactor of the third stage may include ethane and ethylene. In order to separate the ethane and ethylene in the rear stage process, as described above, a complicated process such as a multi-stage distillation column is required, resulting in high operating and investment costs. It may take a lot, so it is necessary to increase the conversion of ethylene to exclude this. Therefore, the present invention was to identify the optimum reaction conditions and catalysts that can maximize the conversion of ethylene in the third step reaction, it is possible to recycle to the first step without having to separate them separately.

The present invention may further include a fourth step of recycling the unreacted gas of the third step to the steam reforming reaction of the first step and the oxidative dimerization reaction of methane after the third step.

In the present invention, the product of the third step may be gas-liquid separated, and the gas phase of the gas-liquid separated second step product may be separated and recycled to the first step. In the gas phase, unreacted methane is the main component in the oxidative dimerization reaction of methane in the first step. The recycled gaseous component to the first stage can be partitioned in oxidative dimerization of methane or steam reforming of methane, or alternatively in proportion to all of the reactions.

When the ratio of oxygen to methane is about 3: 1 in methane oxidative dimerization, about 30% of methane conversion and about 100% conversion are obtained. When the C ₂ yield is about 15 to 20%, the heat of reaction of methane oxidative dimerization reaction Methane required for the steam reforming of methane required to neutralize is about half of the methane involved in methane oxidation dimerization. The conversion rate of methane reforming reaction is about 95%. Therefore, it is necessary to recycle some of the methane recycled to the oxidative dimerization reaction of the methane because it is excessive to recycle all of the unreacted methane to the methane reforming reaction as shown above.

On the other hand, by-products of the oxidative dimerization of methane include CO, H ₂ , H ₂ O and CO ₂ , which are products or reactants of the methane reforming reaction and recycled to the methane reforming reaction rather than the oxidative dimerization reaction of methane. It is advantageous. In conclusion, the amount and composition of the by-products in the oxidative dimerization of methane depends on the reaction conditions, but it is advantageous in terms of the efficiency of the overall process to recycle the remaining by-products to the oxidative dimerization reaction of methane.

After cooling the product from the olefin oligomerization reactor of the third step, it can be separated into the gas phase and liquid phase using a gas-liquid separator. The liquid product consists of C5 + hydrocarbons containing C5 + olefins and water which is the product of the oxidative dimerization reaction of methane. C5 + hydrocarbons can be obtained as gasoline by further hydrogenation processes and converted to light olefins by cracking processes.

In the methane conversion method of the present invention, the conversion rate of methane is 20 to 80%, the selectivity for C 2+ hydrocarbons is 40 to 80%, and the overall yield of C 2+ hydrocarbons may be 15 to 25%.

Figure 2 is a process schematic showing the conversion process of the methane-containing gas in more detail, in the production of C5 + hydrocarbons and synthetic oil or methanol or hydrogen in methane, showing a process according to the invention less energy consumption, lower equipment costs and operating costs . First, the methane-containing reaction gas is introduced into the oxidative dimerization reaction layer 2 of methane in the microchannel reactor 1 through the flow 21, where oxygen is supplied together as the oxidant of the methane. The recycle stream 27 of unreacted and recovered methane is mixed with the stream 21 and fed to the oxidative dimerization reaction bed 2 of methane. After the oxidative dimerization of methane, the product is discharged as stream 22 and quenched by cooling water, where the cooling water is heated to obtain additional high pressure steam. The cooled stream 22 enters the flash column 4, which is a gas-liquid separator, and the phases are separated, and the liquid water is discharged into the stream 24. The water condensed and discharged may be reused as process water after the purification process. The remaining gaseous gas component is a stream 23 which is boosted or adjusted to a pressure of at least 3 atm suitable for the ethylene oligomerization reaction, and then introduced into the olefin oligomerization reactor 5 via heat exchange. After the olefin oligomerization reaction, the product is withdrawn into stream 25 and fed to separator 6 to separate the phases. The gaseous phase separated in separator 6 is recycled to microchannel reactor 1 as stream 27.

On the other hand, the methane-containing reaction gas is introduced into the reforming reaction layer 3 of methane of the microchannel reactor 1 through the stream 31, where steam or carbon dioxide or a mixture thereof is supplied together. At this time, the recycle stream 27 of unreacted and recovered methane is mixed with the stream 31 and supplied to the reforming reaction layer 3 of methane. After methane reforming, the syngas is discharged as stream 32 and quenched by cooling water. The cooled stream 32 is introduced into a flash column, which is a gas-liquid separator, to separate phases, and to separate unreacted liquid water. The remaining gaseous gas component is boosted or adjusted to a pressure suitable for the syngas conversion reaction and then introduced into the syngas conversion device 7. After the syngas shift reaction, the product is discharged to stream 33 and fed to separator 8 to separate the product. Unreacted syngas separated in the separator 8 is recycled to the syngas converter 7 via a flow 34.

Hereinafter, the present invention is described in detail in order to describe the present invention, but the present invention is not limited to the embodiments.

Preparation Example 1: Microchannel Reactor

Inconel 600 was used as a heat resistant material of the microchannel reactor. In the microchannel reactor, Inconel plates are stacked and each layer has a predetermined thickness and they are separated from each other.

The microchannel reactor consists of alternating oxidant dimerization reaction beds of methane and reforming reaction beds of methane. Specifically, the oxidation dimerization reaction layer is three layers, the methane reforming reaction layer is 4 layers, the plate size is 6 cm x 6 cm, the same as the catalyst layer of the oxidation dimerization reaction of methane and the reforming reaction of methane 3 mm thick. Each layer is filled with the respective catalysts and the rear end has a built-in filter to prevent the catalysts from escaping.

(A) to (c) is a schematic diagram showing a microchannel reactor in which the oxidative dimerization reaction of methane and steam reforming reaction of methane produced as described above, and (d) of FIG. Photo of the reactor.

Production example  2 to 6: Microchannel Reactor

A microchannel reactor was manufactured in the same manner as in Preparation Example 1, but the thickness and the number of stacked layers of the catalyst layer were prepared as shown in Table 1.

	SMR Floors / OCM Floors	Plate size (width) x (axial direction)	SMR catalyst layer thickness	OCM catalyst layer thickness	Manufactured reactor
Production Example 1	3rd and 4th floors	6 cm x 6 cm	3 mm	3 mm	3 (d)
Production Example 2	5th / 6th floor	4 cm x 10 cm	3 mm	2 mm	Fig. 5 (a)
Production Example 3	1st floor / 2nd floor	4 cm x 10 cm	4 mm	3 mm	Figure 5 (b)
Production Example 4	3rd and 4th floors	4 cm x 10 cm	4 mm	3 mm	Figure 5 (c)
Production Example 5	3rd and 4th floors	9 cm x 4 cm	3 mm	3 mm	Fig. 5 (d)
Production Example 6	3rd and 4th floors	4 cm x 10 cm	7 mm	4 mm	Figure 5 (e)

In the case of Production Example 5, the inside of the catalyst layer of the oxidative dimerization reaction of methane is coated with a ceramic surface in order to suppress the hydrogen generation by the partial oxidation reaction of methane by the nickel metal component contained in the reactor material during the oxidative dimerization reaction of methane Was prepared.

Production Example  1: of methane Oxidation Dimerization  Preparation of reaction catalyst

Commercial silica (SiO ₂ , Davisil grade 635) was used as a catalyst carrier to prepare a catalyst for oxidative dimerization of methane. First, La, Mn, Na ₂ WO ₄ source La (NO ₃ ) ₃ · 6H ₂ O, Mn (NO ₃ ) ₂ · 4H ₂ O and Na ₂ WO ₄ · 2H ₂ O made by dissolving in distilled water The aqueous solution was supported on the silica carrier by incipient impregnation. The supported catalyst was dried and calcined at 800 ° C. for 5 hours to use as an oxidation dimerization catalyst of methane. The composition of the prepared catalyst is 4.75 Na ₂ WO ₄ /2Mn/0.25La/SiO ₂ , the number indicating the weight percent.

Preparation Example 2 Preparation of Catalyst for Reforming Methane

The catalytic activity was adjusted to prepare a catalyst for the reforming reaction of methane which is suitable for controlling the heat of reaction of the oxidative dimerization reaction of methane.

A catalyst for reforming methane was prepared using gamma-alumina (γ-Al ₂ O ₃ ) as a catalyst carrier. First, Pt, Ni, Mg, and a K source _{(source) Pt (NH 3)} 4 (OH) 2 · xH 2 O (Tetraammineplatinum (II) hydroxide hydrate), Ni (NO 3) 2 · 6H 2 O, Mg ( NO ₃ ) Aqueous solution prepared by dissolving ₂ · 6H ₂ O and KNO ₃ in distilled water was prepared by sequentially supporting the alumina carrier by incipient impregnation. The supported catalyst was dried and calcined at 1100 ° C. for 5 hours to use as a reforming catalyst for methane. The composition of the prepared catalyst is 0.05Pt / 2K / 12Ni-5Mg / Al ₂ O ₃ , the number of which indicates wt%.

Preparation Example 3 Preparation of Catalyst for Reforming Methane

A catalyst for reforming methane having general catalytic activity was prepared. It was prepared in a similar manner to Preparation Example 2, except that it did not carry K and was fired at a lower temperature.

A catalyst for reforming methane was prepared using gamma-alumina (γ-Al ₂ O ₃ ) as a catalyst carrier. First, Pt, Ni and a Mg source of _{Pt (NH 3) 4 (OH} ) 2 · xH 2 O, Ni (NO 3) 2 · 6H 2 O, and Mg (NO _{3) 2} · dissolved 6H ₂ O in distilled water The prepared aqueous solution was prepared by sequentially supporting the alumina carrier by the initial wet impregnation method. The supported catalyst was dried and calcined at 900 ° C. for 5 hours to use as a reforming catalyst for methane. The composition of the prepared catalyst is 0.05Pt / 12Ni-5Mg / Al ₂ O ₃ , the number of which indicates wt%.

Preparation Example 4 Preparation of Catalyst for Ethylene Oligomerization

A zeolite slurry was prepared by mixing 50 g of HZSM-5 (Si / Al = 25, in mole) with a solution containing 13.0 g of phosphoric acid (H ₃ PO ₄ , 85%) and 40 g of distilled water. The slurry was dried at 110 ° C. and calcined at 550 ° C. for 5 hours to prepare a catalyst for ethylene oligomerization.

Production Example  5: fisher Tropsch  Preparation of Cobalt-Based Catalysts for Reaction

30.0 g of gamma-alumina was added to a solution containing 25 mL of distilled water and 26.2 g of cobalt nitrate (Co (NO ₃ ) ₂ 6H ₂ O). The slurry was dried at 100 ° C. for at least 12 hours, and then calcined at 400 ° C. for 5 hours to prepare a 15 wt% Co / Al ₂ O ₃ catalyst.

Preparation Example 6 Preparation of Iron-Based Catalyst for Fischer-Tropsch Reaction

Iron-based Fischer-Tropsch catalysts having a composition of 2.5K / 100Fe / 4Cu / 10Mn / 20Al ₂ O ₃ molar ratio were prepared using mixed-coprecipitation and extrusion-molding.

Preparation Example 7 Preparation of Methanol Synthesis Catalyst

A methanol synthesis catalyst having a composition of 0.75Cu / 1Zn / 0.26Al oxide molar ratio was prepared using mixed-coprecipitation and extrusion-molding.

Example 1

The methane oxidative dimerization catalyst and the methane reforming catalyst prepared in Preparation Examples 1 and 3 above were mounted in the microchannel reactor prepared in Preparation Example 1, and the exothermic and endothermic reactions were thermally neutralized by heat exchange. It was. The volume of the catalyst mounted in the oxidative dimerization reaction layer of methane was 23 cc, and the volume of the catalyst mounted in the reforming reaction layer of methane was 28 cc. The microchannel reactor equipped with the catalyst was heated to 780 ° C by an electric furnace. At this time, the inside of the reactor was allowed to flow at 300cc / min. When the temperature outside the reactor rose to 780 ° C, the gas composition and flow rate of the reforming reaction layer were gradually changed. The amount of gas in the normal reforming reaction is methane 700cc / min, nitrogen 500cc / min (GC standard gas), water 1.4cc / min, and the reaction inlet pressure is 0.4 bar. The amount of gas in the oxidative dimerization reaction of methane was 1500 cc / min of methane, 3000 cc / min of nitrogen (GC standard gas and diluent gas), and 600 cc / min of oxygen, where the reaction inlet pressure was 1.8 bar. The temperature inside the oxidative dimerization reaction layer of methane was 795 ° C., and the temperature inside the reforming reaction layer of methane was 723 ° C. Exothermic reaction of oxidative dimerization of methane was higher than the outer layer temperature of the microchannel reactor. On the other hand, the temperature inside the reforming bed of methane was lower than outside of the microchannel reactor by the endothermic reaction. Methane conversion was calculated based on the internal standard nitrogen. The selectivity of the hydrocarbon product was calculated based on methane.

The results of the oxidative dimerization reaction of methane obtained in Examples 1 to 3 are shown in Table 2 (product distribution of the oxidative dimerization reaction of methane), and the results of the reforming reaction of methane are shown in Table 3 (product distribution of the reforming reaction of methane). ).

Example  2

An oxidation dimerization catalyst of methane was prepared in the same manner as in Preparation Example 1, but a 3Na ₂ WO ₄ /2Mn/0.5La/SiO ₂ catalyst was prepared, and a reforming catalyst of methane was used in the same catalyst as in Preparation Example 2. The gas content of the normal reforming reaction is methane 700cc / min, nitrogen 500cc / min (GC standard gas), water 1.4cc / min, and the reaction inlet pressure is 0.4 bar. The amount of gas in the oxidative dimerization reaction of methane is methane 1300cc / min, nitrogen 2600cc / min (GC standard gas and diluent gas), oxygen 520cc / min, except that the reaction inlet pressure is 1.8 bar and It was done in the same way.

The temperature in the oxidative dimerization reaction layer of methane was 818 ° C, and the temperature in the reforming reaction layer of methane was 710 ° C.

Example  3

As in Example 2, an oxidation dimerization catalyst of methane and a reforming catalyst of methane were used. The volume of the catalyst mounted on the methane oxidative dimerization reaction bed is 20 cc, the volume of the catalyst mounted on the methane reforming reaction bed is 28 cc. , 1.6cc / min of water, where the reaction inlet pressure is 0.5bar. The amount of gas in the oxidative dimerization reaction of methane is methane 1100cc / min, nitrogen 1833cc / min (GC standard gas and diluent gas), oxygen 436cc / min, except that the reaction inlet pressure is 1.1 bar and It was done in the same way.

The temperature inside the oxidative dimerization reaction layer of methane was 782 ° C., and the temperature inside the reforming reaction layer of methane was 778 ° C.

Comparative example  One

The same oxidation dimerization catalyst (3Na ₂ WO _{4 /} 2Mn / 0.5La / SiO ₂ ) of methane was used as in Example ₂ . The catalyst was mounted in an Inconel tube (1/2 inch outer diameter) reactor to conduct oxidative dimerization of methane. Unlike Examples 1 to 3, only the oxidation dimerization reaction of methane was carried out using a single tube. The weight of the catalyst mounted in the oxidative dimerization reaction of methane was 0.2 g. The reactor equipped with the catalyst was heated to 800 ° C. by an electric furnace.

The amount of gas in the oxidative dimerization reaction of methane was methane 166cc / min, nitrogen 139cc / min (GC standard gas and diluent gas), oxygen 55cc / min, and the reaction inlet pressure was 0.2 bar. The temperature inside the oxidative dimerization reaction layer of methane was 850 ° C., but the internal temperature increased due to the exothermic reaction of the oxidative dimerization reaction of methane.

The results of the oxidative dimerization reaction of methane obtained above are shown in Table 2 below. Although the C ₂ yield is similar to that of the Example, only the oxidation dimerization reaction of methane is performed. Unlike the embodiment that simultaneously performs the reforming reaction of methane, if the thermal efficiency is low in the whole process, and the amount of catalyst increases, the flow rate of the total reacting gas increases, the temperature of the reactor rises to 900 ° C. or more, which lowers the C ₂ yield and the stability of the reactor. This can cause serious risks. On the other hand, in the case of using a microchannel reactor having high heat exchange efficiency because the layers for performing oxidative dimerization and methane reforming of methane are alternately arranged as in the embodiment, the reaction heat is easily controlled even when the reactor scale is increased, and the exothermic reaction is divergent. There is an advantage in that heat can be effectively used as a heat source of the endothermic reaction. In a practical example, even if the amount of the catalyst was increased by more than 30 times compared with the comparative example, the temperature of the reactor could be kept constant.

Comparative example  2

0.5 g of a 5Na ₂ WO ₄ /2Mn/0.5La/SiO ₂ catalyst was mounted in a quartz tube (3/4 inch outer diameter) reactor, and the same procedure as in Comparative Example 1 was carried out except that the oxidation dimerization reaction of methane was carried out.

The temperature inside the oxidative dimerization reaction layer of methane was 920 ° C., and the internal temperature greatly increased due to the exothermic reaction of the oxidative dimerization reaction of methane.

The results of the oxidative dimerization reaction of methane obtained above are shown in Table 2 below. Unlike Comparative Example 1, when the amount of catalyst is increased from 0.2g to 0.5g and the reactor material is changed from a metal with good thermal conductivity to a quartz reactor with low thermal conductivity, the temperature inside the reactor is increased from 850 ° C to 920 ° C as in Comparative Example 2. You can see the increase. Therefore, increasing the amount of catalyst further may make it difficult to control the temperature of the reactor and lower the C ₂ yield. On the other hand, combining the oxidative dimerization reaction of methane and the reforming reaction of methane as in the present invention can overcome the problems of temperature control and low thermal efficiency of the process of the conventional adiabatic fixed bed reactor.

product	Example 1	Example 2	Example 3	Comparative Example 1	Comparative Example 2
Methane conversion rate (%)	34.5	32.8	33.5	31.7	33.7
Oxygen conversion rate (%)	98.8	97	93.2	93.2	98.1
Product selectivity (%)
CO	23.1	20.5	26.4	38.1	33.7
CO 2	31.5	30.4	34.2	18.4	32.5
CH 3 CH 3	15.5	18.9	16.2	10.7	15.6
CH 2 = CH 2	29.9	30.2	23.2	32.8	18.2
C 2 total	45.4	49.1	39.4	43.5	33.8
C 2 yield	15.7	16.1	13.2	13.8	11.4

product	Example 1	Example 2	Example 3
Methane conversion rate (%)	95.7	95.6	94.3
Product selectivity
CO	18.0	17.0	18.7
CO 2	5.8	5.4	5.1
H 2	77.0	77.6	76.2
H 2 / CO	4.3	4.6	4.1

Example  4: Oligomerization  reaction

As a catalyst for the olefin (ethylene) oligomerization reaction, the zeolite catalyst prepared in Preparation Example 4 (HZSM-5, Si / Al = 25 mole ratio, supported by 0.7 wt% P) was used. The reaction was carried out in a fixed bed reactor packed with catalyst, and the reactants used in the reaction were simulated gas having a composition similar to that of the product of oxidative dimerization of methane (composition gas composition: nitrogen 5.0%, methane 60.5%). , Ethylene 10.3%, ethane 5.3%, CO 8.0%, CO ₂ 10.9%, volume%). The reaction temperature was 400 ° C., the reaction pressure was 5 bar, and the space velocity (GHSV) was performed at 4,000 h ⁻¹ to convert ethylene into C5 + hydrocarbon. Ethylene conversion was calculated based on the internal standard nitrogen. Selectivity of the hydrocarbon product was calculated on the basis of ethylene. The product distribution of the ethylene oligomerization reaction is shown in Table 4 (product distribution of the ethylene oligomerization reaction).

Example  5

HZSM-5 (Si / Al = 50 mole ratio) catalyst was used, the reaction temperature was 380 ° C., the reaction pressure was 7bar, and the space velocity (GHSV) was the same as that of Example 4, except that the reaction was performed under 4,000 h ⁻¹ . The olefin (ethylene) oligomerization reaction was carried out to convert ethylene into C5 + olefins. The results are shown in Table 4 below.

division	Ethylene Conversion (%)	Product selectivity (%)
		methane	ethane	Propane	Propylene	butane	Butene	C5 +
Example 4	97.1	0.09	One	2.5	One	5.9	5.7	83.8
Example 5	96.6	0.08	0.9	2.3	2.1	5.4	7.8	81.4

Example  6

Experiments were carried out to produce synthetic oils by Fischer-Tropsch reactions in syngas. 0.5 g of the cobalt catalyst prepared in Preparation Example 5 was charged to a 1/2 inch stainless steel fixed bed reactor, and subjected to reduction for 5 hours under an atmosphere of hydrogen (5% by volume H ₂ / He) at 400 ° C. Reaction temperature 230 ℃, reaction pressure 20 kg / cm ² , The molar ratio of reactant carbon monoxide: hydrogen: argon (internal standard) at a rate of 4000 L / kg _cat / hr was fixed at a ratio of 31.5: 63.0: 5.5 and injected into the reactor. The Fischer Tropsch reaction was carried out, and the results of measuring the activity of the catalyst after 40 hours were shown in Table 5 (product distribution of the Fischer-Tropsch reaction).

Example  7

The iron-based Fischer-Tropsch catalyst prepared in Preparation Example 6 was crushed and classified into a size of about 1 mm, 1 g of the catalyst was charged, and charged into a one-stage fixed bed reactor, and activated under a hydrogen atmosphere at atmospheric pressure at 450 ° C. for 12 hours. Reduced. Reaction temperature 280 ℃, reaction pressure 10 kg / cm ² , The molar ratio of reactants carbon monoxide: hydrogen: carbon dioxide: argon (internal standard) at a space velocity of 3600 L / kg _cat / hr was fixed at a ratio of 18.0: 60.5: 16.0: 5.5 and injected into the reactor. Fischer Tropsch reaction was performed, and the results of measuring the activity of the catalyst after the reaction time of 40 hours are shown in Table 5 below.

division	catalyst	CO conversion rate (%)	Hydrocarbon selectivity (%)
			CH 4	C 2 -C 4	C 5 +
Example 6	15% Co / γ-Al 2 O 3	64.4	12.1	11.8	76.1
Example 7	2.5K / 100Fe / 4Cu / 10Mn / 20Al 2 O 3	91.2	6.7	7.5	85.8

Example  8

The methanol synthesis catalyst prepared in Preparation Example 7 was crushed and classified into a size of about 1 mm, 1 g of the catalyst was charged, and charged in a one-stage fixed bed reactor, and reduced by performing an activation process in a hydrogen atmosphere at atmospheric pressure at 280 ° C. for 5 hours. Reaction temperature 250 ℃, reaction pressure 60 kg / cm ² , The composition of gas composition suitable for methanol synthesis reaction is fixed by fixing the molar ratio of carbon monoxide: hydrogen: carbon dioxide: argon (internal standard) as reactants at the space velocity of 4000 L / kg _cat / hr at the ratio of 19.0: 66.5: 9.5: 5.0. H ₂ / (2CO + 3CO ₂ ) = 1 was injected into the reactor to match the gas composition. As a result of the methanol synthesis reaction, a CO conversion of 60.5%, a CO ₂ conversion of 8.1%, and a methanol selectivity of 99.35% were obtained.

Example  9

The methane oxidative dimerization catalyst and the methane reforming catalyst prepared in Preparation Examples 1 and 2 above were mounted in the microchannel reactor prepared in Preparation Example 1, and the exothermic and endothermic reactions were thermally neutralized by heat exchange. The flow of reactants is in the same direction. The microchannel reactor equipped with the catalyst was heated to 760 ° C by an electric furnace. At this time, the inside of the reactor was allowed to flow at 300 cc / min. When the temperature outside the reactor rose to 760 ° C, the gas composition and flow rate of the reforming reaction layer were gradually changed. The gas content of the normal reforming reaction is methane 300 cc / min, hydrogen 300 cc / min, nitrogen 400 cc / min (GC standard gas), water 0.24 cc / min, and the reaction inlet pressure is 0.5 barg. The amount of gas in the oxidative dimerization reaction of methane was 1300 cc / min of methane, 2000 cc / min of nitrogen (GC standard gas and diluent gas), and 520 cc / min of oxygen, where the reaction inlet pressure was 1.4 barg. The temperature inside the oxidative dimerization reaction layer of methane was 815 ° C. The results of the oxidative dimerization reaction of methane and the reforming reaction of methane obtained above are shown in Table 6 below.

Example  10

The methane oxidative dimerization catalyst and the methane reforming catalyst prepared in Preparation Examples 1 and 2 were mounted in the microchannel reactor prepared in Preparation Example 2, and the exothermic and endothermic reactions were thermally neutralized by heat exchange. The flow of reactants is in the same direction. The microchannel reactor equipped with the catalyst was heated to 760 ° C by an electric furnace. At this time, the inside of the reactor was allowed to flow at 300cc / min. When the temperature outside the reactor rose to 760 ° C, the gas composition and flow rate of the reforming reaction layer were gradually changed. The gas content of the normal reforming reaction is methane 700 cc / min, hydrogen 600 cc / min, nitrogen 700 cc / min (GC standard gas), water 0.56 cc / min, and the reaction inlet pressure is 0.8 barg. The amount of gas in the oxidative dimerization reaction of methane was 2500 cc / min of methane, 3000 cc / min of nitrogen (GC standard gas and diluent gas), and 1000 cc / min of oxygen, with a reaction inlet pressure of 1.74 barg. The temperature inside the oxidative dimerization reaction layer of methane was 829 ° C.

Example  11

The methane oxidative dimerization catalyst and the methane reforming catalyst prepared in Preparation Examples 1 and 2 were mounted in the microchannel reactor prepared in Preparation Example 3, and the exothermic and endothermic reactions were thermally neutralized by heat exchange. The flow of reactants is in the same direction. The microchannel reactor equipped with the catalyst was heated to 760 ° C by an electric furnace. At this time, the inside of the reactor was allowed to flow at 300 cc / min. When the temperature outside the reactor rose to 760 ° C, the gas composition and flow rate of the reforming reaction layer were gradually changed. The gas content of the normal reforming reaction is 200 cc / min of methane, 200 cc / min of hydrogen, 300 cc / min of nitrogen (GC standard gas), and 0.16 cc / min of water, where the reaction inlet pressure is 0.3 barg. The amount of gas in the oxidative dimerization reaction of methane was 600 cc / min of methane, 700 cc / min of nitrogen (GC standard gas and diluent gas), and 240 cc / min of oxygen, and the reaction inlet pressure was 0.9 barg. The temperature inside the oxidative dimerization reaction layer of methane was 813 ° C.

Example  12

The methane oxidative dimerization catalyst and the methane reforming catalyst prepared in Preparation Examples 1 and 2 were mounted in the microchannel reactor prepared in Preparation Example 4, and the exothermic and endothermic reactions were thermally neutralized by heat exchange. The flow of reactants is in the same direction. The microchannel reactor equipped with the catalyst was heated to 760 ° C by an electric furnace. At this time, the inside of the reactor was allowed to flow at 300 cc / min. When the temperature outside the reactor rose to 760 ° C, the gas composition and flow rate of the reforming reaction layer were gradually changed. The gas content of the normal reforming reaction is methane 350 cc / min, hydrogen 300 cc / min, nitrogen 400 cc / min (GC standard gas), carbon dioxide 180 cc / min, water 0.18 cc / min, and the reaction inlet pressure is 0.5 barg. The amount of gas in the oxidative dimerization reaction of methane was 1500 cc / min of methane, 2500 cc / min of nitrogen (GC standard gas and diluent gas), and 600 cc / min of oxygen, and the reaction inlet pressure was 0.9 barg. The temperature inside the oxidative dimerization reaction layer of methane was 822 ° C.

Example  13

The methane oxidative dimerization catalyst and the methane reforming catalyst prepared in Preparation Examples 1 and 2 were mounted in the microchannel reactor prepared in Preparation Example 5, and the exothermic and endothermic reactions were thermally neutralized by heat exchange. The flow of reactants is in the same direction. The microchannel reactor equipped with the catalyst was heated to 760 ° C by an electric furnace. At this time, the inside of the reactor was allowed to flow at 300 cc / min. When the temperature outside the reactor rose to 760 ° C, the gas composition and flow rate of the reforming reaction layer were gradually changed. The gas content of the normal reforming reaction is methane 300 cc / min, hydrogen 300 cc / min, nitrogen 300 cc / min (GC standard gas), water 0.24 cc / min, and the reaction inlet pressure is 0.4 barg. The amount of gas in the oxidative dimerization reaction of methane was methane 1400 cc / min, nitrogen 1800 cc / min (GC standard gas and diluent gas), oxygen 560 cc / min, and the reaction inlet pressure was 0.9 barg. The temperature inside the oxidative dimerization reaction layer of methane was 829 ° C.

Example  14

The methane oxidative dimerization catalyst and the methane reforming catalyst prepared in Preparation Examples 1 and 2 were mounted in the microchannel reactor prepared in Preparation Example 2, and the exothermic and endothermic reactions were thermally neutralized by heat exchange. The flow of reactants is in the same direction. The microchannel reactor equipped with the catalyst was heated to 760 ° C by an electric furnace. At this time, the inside of the reactor was allowed to flow at 300 cc / min. When the temperature outside the reactor rose to 760 ° C, the gas composition and flow rate of the reforming reaction layer were gradually changed. The gas reforming rate is 800 cc / min for methane, 600 cc / min for hydrogen, 700 cc / min for nitrogen (GC standard gas), and 0.64 cc / min for water. The reaction was carried out by raising the pressure to 2.1 barg. The amount of gas in the oxidative dimerization reaction of methane was 2500 cc / min of methane, 3000 cc / min of nitrogen (GC standard gas and diluent gas), and 1000 cc / min of oxygen, where the reaction inlet pressure was 1.7 barg. The temperature inside the oxidative dimerization reaction layer of methane was 833.

Comparative example  3

The methane oxidative dimerization catalyst and the methane reforming catalyst prepared in Preparation Examples 1 and 2 were mounted in the microchannel reactor prepared in Preparation Example 6, and the exothermic and endothermic reactions were thermally neutralized by heat exchange. The flow of reactants is in the same direction. The microchannel reactor equipped with the catalyst was heated to 760 ° C by an electric furnace. At this time, the inside of the reactor was allowed to flow at 300 cc / min. When the temperature outside the reactor rose to 760 ° C, the gas composition and flow rate of the reforming reaction layer were gradually changed. The gas content of the normal reforming reaction is methane 900 cc / min, hydrogen 600 cc / min, nitrogen 700 cc / min (GC standard gas), water 0.72 cc / min, and the reaction inlet pressure is 0.7 barg. The amount of gas in the oxidative dimerization reaction of methane was methane 2700 cc / min, nitrogen 3300 cc / min (GC standard gas and diluent gas), oxygen 1080 cc / min, and the reaction inlet pressure was 1.8 barg. The temperature inside the oxidative dimerization reaction layer of methane was 835 ° C.

Comparative example  4

The methane oxidative dimerization catalyst and the methane reforming catalyst prepared in Preparation Examples 1 and 2 were mounted in the microchannel reactor prepared in Preparation Example 2, and the exothermic and endothermic reactions were thermally neutralized by heat exchange. The flow of reactants is in the same direction. The microchannel reactor equipped with the catalyst was heated to 760 ° C by an electric furnace. At this time, the inside of the reactor was allowed to flow at 300 cc / min. When the temperature outside the reactor rose to 760 ° C, the gas composition and flow rate of the reforming reaction layer were gradually changed. The gas content of the normal reforming reaction is 700 cc / min of methane, 700 cc / min of nitrogen (GC standard gas), and 1.69 cc / min of water, where the reaction inlet pressure is 0.95 barg. The amount of gas in the oxidative dimerization reaction of methane was 2500 cc / min of methane, 3000 cc / min of nitrogen (GC standard gas and diluent gas), and 1000 cc / min of oxygen, where the reaction inlet pressure was 1.7 barg. The temperature inside the oxidative dimerization reaction layer of methane was 830 ° C.

Comparative example  5

The methane oxidative dimerization catalyst and the methane reforming catalyst prepared in Preparation Examples 1 and 3 above were mounted in the microchannel reactor manufactured in Preparation Example 2, and the exothermic and endothermic reactions were thermally neutralized by heat exchange. The flow of reactants is in the same direction. The microchannel reactor equipped with the catalyst was heated to 760 ° C by an electric furnace. At this time, the inside of the reactor was allowed to flow at 300 cc / min. When the temperature outside the reactor rose to 760 ° C, the gas composition and flow rate of the reforming reaction layer were gradually changed. The gas content of the normal reforming reaction is 700 cc / min of methane, 700 cc / min of nitrogen (GC standard gas), and 1.13 cc / min of water, where the reaction inlet pressure is 0.75 barg. The amount of gas in the oxidative dimerization reaction of methane was 2500 cc / min of methane, 3000 cc / min of nitrogen (GC standard gas and diluent gas), and 1000 cc / min of oxygen, where the reaction inlet pressure was 1.7 barg. The temperature inside the oxidative dimerization reaction layer of methane was 825 ° C.

Comparative example  6

Example 10 was carried out in the same manner as in Example 10, except that the fluid flow of the reactants of the methane oxidative dimerization reaction and the methane reforming reaction was reversed.

	Oxidation Dimerization of Methane								Reforming
	Conversion rate		Product selectivity (%)					C2 yield (%)	CH 4 conversion (%)
	methane(%)	Oxygen(%)	CO	CO 2	ethane	Ethylene	C2 toal
Example 9	33.7	97.5	21.4	25.8	16.2	36.6	52.8	17.8	91.2
Example 10	34.3	96.8	22.3	24.9	15.9	36.9	52.8	18.1	90.8
Example 11	33.1	94.3	25.3	24.9	18.1	31.7	49.8	16.5	90.3
Example 12	35.40	95.10	23.7	31.4	17.2	27.7	44.9	15.9	91.1
Example 13	34.9	96.3	24.6	28.1	14.9	32.4	47.3	16.5	90.7
Example 14	36.1	95.8	22.7	32.1	15.6	29.6	45.2	16.3	75.4
Comparative Example 3	32.4	92.3	37.4	26.8	14.3	21.5	35.8	11.6	93.2
Comparative Example 4	33.2	94.8	35.3	20.7	16.3	27.7	44.0	14.6	98.5
Comparative Example 5	34.5	93.9	33.2	23.6	14.8	28.4	43.2	14.9	97.2
Comparative Example 6	35.1	93.6	23.6	37.4	13.6	25.4	39.0	13.7	95.3

As carried out in Examples 9 to 14 of the present invention, the catalyst layer thickness of the oxidative dimerization reaction of methane is in an appropriate range, or a catalyst having a low catalytic activity in the reforming reaction of methane is used, and the molar ratio of (water vapor + carbon dioxide) / methane Lowering to near 1 prevents the methane conversion of the methane reforming reaction from being rapidly converted at low temperatures, which is more effective in controlling the heat of reaction of the oxidative dimerization reaction of methane by endothermic reaction, which is advantageous for obtaining a high product, and more easily in the reactor. Could drive.

On the other hand, when the catalyst layer of the oxidative dimerization reaction of methane is too thick (Comparative Example 3), it is difficult to control the reaction heat of the oxidative dimerization reaction of methane, and the product yield is low, and when the molar ratio of water vapor / methane is kept to 3, the methane conversion is performed at low temperature. It was difficult to control the temperature of the oxidative dimerization reaction of methane (Comparative Example 4). In addition, the high catalytic activity of the methane reforming reaction and the high molar ratio of water vapor / methane showed similar yields even when the methane conversion of the methane reforming reaction was high (Comparative Example 5). As shown in Comparative Example 6, when the fluid flow of the oxidative dimerization reaction of methane and the fluid flow of the reforming reaction of methane are reversed, the product yield is lowered due to the high temperature in the upper catalyst layer of the oxidative dimerization reaction of methane.

Claims (15)

1. With an exothermic reaction flow path and an endothermic reaction flow path, the temperature (T ₁ ) of the exothermic reaction flow path where the oxidative dimerization reaction of methane is performed is higher than the temperature (T ₂ ) of the endothermic reaction flow path where the steam reforming reaction of methane is performed. In a heat exchange reactor in which heat is transferred from an exothermic reaction passage to an endothermic reaction passage, methane characterized in that it comprises a first step of performing oxidative dimerization reaction of methane in the exothermic reaction passage and performing steam reforming reaction of methane in the endothermic reaction passage. How to switch.
2. The method of claim 1, wherein in the first step, ethylene and / or ethane are formed as products as C2 hydrocarbons by oxidative dimerization of methane, and synthesis gas is formed as products by steam reforming of methane,

   Optionally, a second step of preparing a synthetic oil from methanol synthesis, hydrogen production, ammonia production or Fischer-Tropsch reaction from the synthesis gas formed in the first step;

   A third step of preparing a liquid hydrocarbon product by ethylene oligomerization reaction from the C2 hydrocarbon formed in the first step; And

   And a fourth step of recycling the unreacted gas of the third step to the steam reforming reaction of the first step and the oxidative dimerization reaction of the methane.
3. One or more exothermic reaction flow passages and two or more endothermic reaction flow passages are alternately provided, and the heat exchanger is heat transfer from the exothermic reaction flow path to the endothermic reaction flow path because the temperature (T ₁ ) of the exothermic reaction flow path is higher than the temperature (T ₂ ) of the endothermic reaction flow path. In a microchannel reactor,

   The exothermic reaction channel is filled with a catalyst for oxidative dimerization of methane,

   The endothermic reaction flow path is filled with an endothermic catalyst,

   In the exothermic reaction passage through heat exchange with the endothermic reaction passage, the reaction temperature is controlled within the range of 800 ℃ ± 50 ℃,

   A heat exchange microchannel reactor, characterized in that the reaction temperature in the endothermic reaction passage through the heat exchange with the exothermic reaction passage is controlled within the range of 750 ℃ ± 50 ℃.
4. One or more exothermic reaction flow passages and two or more endothermic reaction flow passages are alternately provided, and the heat exchanger is heat transfer from the exothermic reaction flow path to the endothermic reaction flow path because the temperature (T ₁ ) of the exothermic reaction flow path is higher than the temperature (T ₂ ) of the endothermic reaction flow path. In a microchannel reactor,

   The exothermic reaction channel is filled with a catalyst for oxidative dimerization of methane,

   The endothermic reaction flow path is filled with an endothermic catalyst,

   A heat exchange microchannel reactor, characterized in that the thickness of the catalyst bed in the exothermic reaction passage located between two adjacent endothermic reaction passages to control the calorific value in the exothermic reaction passage is controlled within a range of 1 to 5 mm.
5. The method of claim 4, wherein the thickness of the catalyst bed in the endothermic reaction passage is controlled to be within a range of 0.1 to 2 times the thickness of the catalyst layer in the exothermic reaction passage so as to remove the calorific value downstream in the exothermic reaction passage. Heat exchange microchannel reactor.
6. One or more exothermic reaction flow passages and two or more endothermic reaction flow passages are alternately provided, and the heat exchanger is heat transfer from the exothermic reaction flow path to the endothermic reaction flow path because the temperature (T ₁ ) of the exothermic reaction flow path is higher than the temperature (T ₂ ) of the endothermic reaction flow path. In a microchannel reactor,

   The exothermic reaction channel is filled with a catalyst for oxidative dimerization of methane,

   The endothermic flow path is filled with a catalyst for the reforming reaction of methane,

   The reforming reaction of the methane in the endothermic reaction flow path is characterized in that the conversion of methane compared to the equilibrium conversion in the reaction conditions to 95% or less by controlling the catalytic activity.
7. One or more exothermic reaction flow passages and two or more endothermic reaction flow passages are alternately provided, and the heat exchanger is heat transfer from the exothermic reaction flow path to the endothermic reaction flow path because the temperature (T ₁ ) of the exothermic reaction flow path is higher than the temperature (T ₂ ) of the endothermic reaction flow path. In a microchannel reactor,

   The exothermic reaction channel is filled with a catalyst for oxidative dimerization of methane,

   The endothermic flow path is filled with a catalyst for the reforming reaction of methane,

   Methane conversion rate in methane reforming in the endothermic flow path to lower the rate of oxidative dimerization of methane upstream in the exothermic flow path to suppress rapid temperature increase and to remove the calorific value of the oxidative dimerization reaction of methane downstream in the exothermic flow path. Heat exchange microchannel reactor, characterized in that controlled in the range of 60% to 95%.
8. 8. The heat exchange microchannel reactor according to any one of claims 3 to 7, wherein the fluid flow in the exothermic reaction flow path and the fluid flow in the endothermic reaction flow path are in the same direction.
9. 8. The heat exchange microchannel reactor according to claim 6 or 7, wherein a part of the product containing the unreacted methane of the oxidative dimerization reaction of methane is recycled to the reactant in the endothermic flow path for performing the reforming of methane.
10. The heat exchange microchannel according to any one of claims 3 to 5, wherein the catalyst packed in the endothermic reaction flow path is a catalyst for dehydrogenation of ethane, and ethylene is produced through dehydrogenation of ethane in the endothermic reaction flow path. Reactor.
11. The heat exchange microchannel reactor according to any one of claims 3 to 5, wherein the endothermic reaction is a catalyst-free cracking reaction of ethane and ethylene is produced through the cracking reaction of ethane in the endothermic reaction passage.
12. The process of claim 1 or 2, wherein the first step is carried out in the heat exchange microchannel reactor of any one of claims 3-9.
13. A process for producing a gas product by converting methane or ethane in the heat exchange microchannel reactor of claim 3.
14. The method for producing a gas product according to claim 13, wherein an oxidation dimerization reaction of methane is carried out in an exothermic passage to convert C2 or more hydrocarbons from a methane-containing gas into C2 or more hydrocarbons including ethylene and / or ethane.
15. The method of claim 13, wherein the synthesis gas from methane-containing gas is produced by performing reforming reaction of methane in the endothermic reaction flow path.

    A method for producing a gas product, characterized in that ethylene is produced from ethane by carrying out dehydrogenation of ethane or cracking of ethane in an endothermic reaction passage.

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