KR20180077399A - Highly efficient conversion method of methane via combined exothermic and endothermic reaction - Google Patents

Abstract

The present invention relates to a conversion method of methane, which comprises a first step of including an exothermic reaction flow path and an endothermic reaction flow path to perform oxidative coupling of methane in the exothermic reaction flow path, and to perform a steam reforming reaction of methane in the endothermic reaction flow path, in a heat exchange reactor in which heat is transferred from the exothermic reaction flow path to the endothermic reaction flow path since a temperature (T1) of the exothermic reaction flow path that the oxidative coupling of methane is performed, is higher than a temperature (T2) of the endothermic reaction flow path that the steam reforming reaction of methane is performed.

Description

BACKGROUND OF THE INVENTION 1. Field of the Invention The present invention relates to a method for converting methane to methane by an exothermic reaction and an endothermic reaction,

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

Methane, which is the main component of natural gas, is mainly used as a heat source for power generation and household use, but some of them are used as raw materials for producing liquid fuels and chemicals. The methane conversion reaction for this purpose is mainly a reforming reaction of methane, which is a synthesis gas production process. These syngas are used as raw materials for methanol synthesis, hydrogen production, ammonia production, or synthetic oil production process by Fischer-Tropsch reaction.

On the other hand, the oxidative coupling of methane is a reaction in which methane and oxygen react on the catalyst to convert into ethylene and ethane, and various side reactions are accompanied. Methane and oxygen react with each other to produce ethylene or ethane according to the same reaction as in the reaction scheme 1) or the reaction scheme 2). In this process, a very strong exothermic reaction may occur with side reactions such as reaction scheme 3) and reaction scheme 4) . If the reaction temperature is increased due to the inability to regulate the reaction heat of the reaction, the reaction proceeds more with CO and CO ₂ , and the yield of the desired C ₂ hydrocarbon compound is lowered. Therefore, the control of the heat of reaction in the quantitative oxidation reaction of methane is important for the yield of the product as well as the stability of the reaction process.

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,? H ₂₉₈ = -518.7 kJ / mol 3)

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

The reforming reaction of methane is strong endothermic reaction as shown in the following reaction, unlike the oxidation reaction of methane. That is, a large amount of heat must be supplied from the outside in order for the methane reforming reaction to take place.

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

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

Methane is a very stable compound, and when it is converted into another compound, it has very strong endothermic reaction or exothermic reaction, which causes a problem of low thermal efficiency. Studies on each of the above reactions have been conducted so far, but there are very few cases in which two reactions are implemented in the same reactor.

The present inventors have realized that the reaction in which the whole reaction is thermally neutralized by using a reactor having excellent heat exchange in the reactor in which the two reactions are carried out is realized and the products of these reactions are recycled to the respective reactions of the unreacted products through the secondary reaction, And the present invention has been completed.

It is an object of the present invention to provide a highly efficient methane conversion method by improving the thermal efficiency of the reaction and integrating with the secondary reaction by allowing the reaction to be carried out under thermal neutral conditions in the conversion of methane to other compounds.

In order to achieve the above object, a first aspect of the present invention is a method for reforming methane comprising the steps of: providing an exothermic reaction channel and an endothermic reaction channel, wherein a temperature (T ₁ ) of an exothermic reaction channel, Wherein the oxidation reaction of the methane is performed in an exothermic reaction channel in a heat exchange reactor in which heat transfer from the exothermic reaction channel to the endothermic reaction channel becomes higher than the temperature (T ₂ ) of the endothermic reaction channel being performed, And a first step carried out in the flow path.

Hereinafter, the present invention will be described in more detail.

The methane conversion method of the present invention is characterized in that the thermal neutralization reaction is performed by carrying out the oxidation quantification reaction of methane and the steam reforming reaction of methane in one reactor to exchange heat of reaction.

To this end, the methane conversion method of the present invention is characterized in that the oxidation quantification reaction of methane and the steam reforming reaction of methane are carried out by providing an exothermic reaction path and an endothermic reaction path so that the temperature of the exothermic reaction path ₁ ) is higher than the temperature (T ₂ ) of the endothermic reaction channel in which the steam reforming reaction of methane is carried out, thereby performing heat transfer from the exothermic reaction channel to the endothermic reaction channel. Preferably, the exothermic reaction channel in which the oxidation reaction of methane is performed and the endothermic reaction channel in which the steam reforming reaction of methane are performed may be disposed adjacent to each other.

In one 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 quantitative oxidation of methane or a catalyst for steam reforming reaction of methane.

The methane conversion method of the present invention is characterized in that in the heat exchange reactor, the oxidation quantification reaction of methane is performed in an exothermic reaction channel and the steam reforming reaction of methane is performed in an endothermic reaction channel. At this time, each reaction occurs while passing through the channel catalyst layer of the heat exchange reactor.

The reactant to be injected into the heat-exchange reactor of the first stage contains methane as a main component and does not have to be 100% methane. It can be derived from natural gas or petrochemical by-products. Further, a small amount of ethane, propane, . ≪ / RTI >

The product of the oxidation quantification reaction of methane may comprise ethane, ethylene, CO, CO ₂ , H ₂ , H ₂ O, trace C3 + hydrocarbons and unreacted methane. Preferably, said first step may be a step wherein ethylene and / or ethane as C ₂ hydrocarbon is formed into the product by the oxidative quantification reaction of methane. The production ratio of ethane and ethylene is similar, but the rate of ethylene production tends to increase as the reaction temperature increases. Oxidation of methane is a strong exothermic reaction. If the heat of reaction is not controlled well, 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.

In the present invention, 5 wt% NaWO ₄ , 2 wt% Mn, and 0.3 wt% La / SiO ₂ catalyst are used as catalysts for the oxidation reaction of methane. However, the present invention is not limited thereto.

In the present invention, it is preferable that the reaction for oxidizing methane is performed at a reaction temperature of 700 to 900 ° C, a reaction pressure of 1 to 10 bar, a methane / oxygen molar ratio of 2 to 10, and a space velocity of 1000 to 50000 h ^-1 .

The steam reforming reaction of another methane in the first step is a reaction in which methane reacts with water vapor, carbon dioxide, or a mixture thereof to produce a syngas (CO + H ₂ ). The steam reforming reaction of methane occurs at 700 ~ 900 ℃, which is similar to the oxidation reaction of methane oxidation, 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 steam / methane of 1 to 4. Since the ratio of H ₂ / CO of the synthesis gas obtained in the steam reforming reaction of methane is as high as 3 or more, in the present invention, the subsequent reaction utilizing these synthesis gases is a methanol synthesis reaction or a Fischer-Tropsch reaction In order to lower the H ₂ / CO ratio, the H ₂ O / CH ₄ ratio was lowered and carbon dioxide was added to some of the reactions. For example, when the molar ratio of CH ₄ : H ₂ O: CO ₂ is about 1: 1.5 to 2.0: 0.4 to 0.8 and the reaction temperature is 800 to 850 ° C, the H ₂ / CO molar ratio is about 2.0 to 2.5, A synthesis gas suitable for the Fischer-Tropsch reaction can be obtained.

Nickel-based catalysts can be used as catalysts for the steam reforming reaction of methane. In the present invention, a 12 wt% Ni / Al ₂ O ₃ catalyst is used, but the present invention is not limited thereto.

In the present invention, it is important that the oxidation reaction of methane and the steam reforming reaction of methane are in close contact with each other in the same reactor to transfer reaction heat. For this, a microchannel reactor was used in the present invention.

The microchannel reactor has a structure in which plates are vertically stacked at regular intervals, and each reaction is alternated in a plate channel, and each channel is filled with a catalyst. That is, the steam reforming reaction layer of methane exists in the lower part and the upper part of the oxidation oxidation reaction layer of methane, and heat transfer can be performed by passing the reaction heat up and down each reaction layer. Therefore, the reaction layer may be an exothermic reaction channel or an endothermic reaction channel. The microchannel reactor can be suitable for reactions with a large heat transfer area compared to the volume of the reactor and having a high reaction heat.

In the microchannel reactor, the temperature control of the oxidation-reduction reaction layer of methane can be controlled by controlling the injection amount of the reactant in the steam reforming reaction of methane to the oxidation-reduction reaction material. When the reactant of the steam reforming reaction of methane increases, the endothermic reaction may increase, and the temperature of the oxidation reaction of methane may decrease. Combining the quantification reaction of methane with the steam reforming reaction of methane as in the present invention has an advantage in that the waste heat of the quantification reaction of methane can be utilized as the heat of the reforming reaction of methane and the reaction heat of the methanation reaction And it is possible to solve the problems of the prior art in which much higher temperature steam must be used in order to control the fever of the high temperature reaction.

Since the product of the oxidative quantification reaction of methane in the first step has a high temperature of 700 ° C or more, it needs to be rapidly cooled to a certain level or less. This is because the second reaction can be performed when the highly reactive ethylene product is maintained at a high temperature for a long time. The cooling of the high temperature product can be achieved by cooling water or steam. The high-pressure or high-temperature steam generated in this process can be utilized as a utility steam for the process. The cooled reaction product is further cooled and then supplied to a gas-liquid separator to carry out gas-liquid separation to separate and recover the liquid product. The gas product is preheated, and then the olefin oligomer / RTI >

Alternatively, all the products of the first stage may be introduced into the olefin oligomerization reactor in the subsequent stage without separately performing the secondary cooling and gas-liquid separation. Wherein the temperature of the product stream after the first cooling is controlled in the range of 250 DEG C to 500 DEG C which is the temperature of the ethylene oligomerization reaction. When the reaction product of the first stage is cooled and then directly fed into the ethylene oligomerization reactor without secondary cooling, the process is simplified by omitting the secondary cooler, the gas-liquid separator and the preheating device, and no separate preheating is required. While there is an advantage that the thermal efficiency can be improved, there may be a water effect on the catalyst of the downstream ethylene conversion reactor.

The present invention may further comprise, after the first step, a second step of producing synthetic oil from the synthesis gas formed in the first step by methanol synthesis, hydrogen production, ammonia production, or Fischer-Tropsch reaction have.

The second step is a step of converting the synthesis gas (CO + H ₂ ) which is the steam reforming reaction product of methane in the first step. There are various types of syngas conversion as follows. Production of wax or synthetic oil from synthetic gas by Fischer-Tropsch reaction, production of methanol from syngas, and production of hydrogen by a water gas conversion reaction of syngas. The present invention is not particularly limited to this. Depending on the kind of the syngas conversion process, the synthesis gas pretreatment process may be changed after the steam reforming process.

For example, if the syngas conversion process is a Fischer-Tropsch process, it is necessary to quench the high temperature syngas produced by the steam reforming reaction and to remove the liquid water through the gas-liquid separator. Since the Fischer-Tropsch reaction is generally a high pressure reaction of 10 bar or more, it is necessary to increase the pressure of the synthesis gas before the Fischer-Tropsch reaction using a compressor. On the other hand, the Fischer-Tropsch reaction requires an appropriate H ₂ / CO ratio of the synthesis gas depending on the catalyst used.

In the case of the cobalt-based catalyst, H _{2 /} CO = 1.8-2.2 is preferable, and if it is lower than the above range, coke is formed on the catalyst or the reaction rate is slowed, and if it is higher than the above range, the production rate of the light hydrocarbon is increased. On the other hand, the iron-based catalyst has a relatively less influence on the H ₂ / CO ratio, but is preferably about 0.7 to 3. The H ₂ / CO ratio can be adjusted by the composition of the reactants in the steam reforming reaction. For example, to adjust the H ₂ / CO ratio to about 2.1 suitable for cobalt-based catalysts, the amount of steam may be adjusted or CO ₂ may be introduced. It is also possible to control the H ₂ / CO ratio by post-treatment at the end of the steam reforming reaction. For example, the H ₂ / CO ratio can be controlled by decreasing H ₂ by a separation membrane or PSA (pressure swing adsorption) method.

In the cobalt-based catalyst, the Fischer-Tropsch reaction may be carried out at 190 to 240 ° C., 10 to 30 bar, SV = 1000 to 4000 and the gas composition described above, and the reactor may be a stationary phase, slurry or microchannel reactor. The Fischer-Tropsch reaction in the iron-based catalyst may be carried out at a temperature of 250 to 350 ° C, 10 to 30 bar, and SV = 500 to 3000 in the above gas composition, and the reactor may be a stationary phase, a slurry reactor or a circulating fluidized bed reactor.

When the syngas conversion process is a methanol synthesis process, it is necessary to rapidly cool the high-temperature syngas produced by the steam reforming reaction and remove the liquid water by gas-liquid separation thereof, and then raise the pressure suitable for methanol synthesis. The methanol synthesis reaction is carried out at a reaction temperature of 230 ~ 280 ° C, 10 ~ 110 bar, SV = 1000 ~ 5000 and a synthesis gas containing part of CO ₂ in the syngas and H ₂ / (2CO + 3CO ₂ ) = 1 Gas composition. Cu / ZnO / Al ₂ O ₃ is widely used as a catalyst, and a fixed bed or a slurry reactor can be used as the reactor.

When the synthesis gas conversion process is a hydrogen synthesis process, the high-temperature synthesis gas produced by the steam reforming reaction is cooled to a temperature suitable for the water gas conversion reaction by rapid cooling, and then introduced into the water gas conversion reactor. At this time, since the water gas conversion reaction requires water as a reactant, there is no need to remove the water separately and additional water can be supplied as needed. Depending on the concentration of CO in the syngas, the water gas conversion reaction can be composed of one or two stages. When the concentration is high, most of the CO is converted to hydrogen through the high temperature conversion reaction and the CO concentration is further lowered at the low temperature reaction. As a method of separating the generated hydrogen, PSA (pressure swing adsorption) or a separation membrane can be used.

The present invention may further comprise, after the second step, a third step of producing a liquid hydrocarbon product by an oligomerization reaction from C ₂ hydrocarbons formed in the first step.

The third step is an ethylene conversion reaction in which the C ₂ hydrocarbons formed in the first step are oligomerized and converted to a hydrocarbon having a high number of carbon atoms (C 5 +), which can be mainly present in a liquid state. In this process, C ₃ -C ₄ hydrocarbons can also be partially produced as byproducts.

In the ethylene oligomerization reaction, paraffinic hydrocarbons and aromatics can be produced according to the severity of the reaction. That is, as the reaction temperature, the reaction pressure, and the acidity of the catalyst increase, the production rate of paraffinic hydrocarbons or aromatics may be higher than that of olefinic hydrocarbons, and thus the selectivity of olefins may be controlled .

The ethylene oligomerization reaction may be carried out on a solid acid catalyst.

The solid acid catalyst is not particularly limited as long as it is in the form of a solid having a Bronsted acid point, and examples thereof include clay minerals such as kaolin; A solid catalyst in which a carrier such as a clay mineral is impregnated and supported with an acid such as sulfuric acid or phosphoric acid; Acid form ion exchange resins; Aluminum phosphates; Mesoporous silica alumina such as Al-MCM41; Zeolites; Or ITO-2 or a layered zeolite such as ITQ-2, and a solid acid catalyst commonly used in the art can be used. The solid acid catalyst preferably has a large surface area and more preferably has a large acid point. Among these solid acid catalysts, it is preferable that the acid strength is not too high, for example, zeolite.

The zeolite catalyst is preferably a zeolite having a median pore diameter of 0.45 nm to 0.55 nm. In the case of zeolites with pores smaller than 0.45 nm, the production of C5 + products becomes difficult and it becomes difficult for the products to escape from the pores. Further, in the case of a zeolite having pores larger than 0.55 nm, a large amount of coke is produced in the zeolite pores, the catalyst life is shortened, and the yield of heavy oil is increased.

The zeolite used as the catalyst in the third step preferably has a ratio of Si / Al of 10 to 100. If the Si / Al ratio is less than 10, the acid point of the zeolite is excessive, so that the production of coke increases, the catalyst life is shortened, and the thermal and hydrothermal stability of the zeolite may be deteriorated. When the Si / Al ratio of the zeolite is larger than 100, the acid site is too small to reduce the reactivity of the zeolite catalyst.

ZSM-5, ZSM-11, ZSM-12, ZSM-35 and MCM-22 are preferred zeolites for use in the catalyst of the third step. Most preferably, ZSM-5 and MCM-22 can be used. The zeolite used as the catalyst may be one in which a transition metal or a basic metal ion is introduced to control the acid property, and La or P may be introduced to control acid characteristics and hydrothermal stability. Further, Ga, Zn and Pt may be introduced into the zeolite to promote aromatization. As another form of the oligomerization catalyst, a mesoporous silica-alumina catalyst containing Ni can be used.

In order to actually apply the zeolite catalyst to the reactor, molding is required, and silica, alumina, silica-alumina, clay or the like can be used as a binder for the zeolite for molding.

The third step may be carried out in an ethylene oligomerization reactor (second reactor). The ethylene oligomerization reactor used in the present invention is not limited in form, and may be a fixed bed reactor, a fluidized bed reactor, a mobile phase reactor, or the like.

In the third step, a known reaction process using a fixed bed reactor, a fluidized bed reactor, a mobile phase reactor or the like may be applied. The third step may be carried out in a batch, semi-continuous or continuous manner, but is preferably carried out continuously. Further, a single reactor may be used, but a plurality of Reactor can be used. However, since the reaction in the third step is an exothermic reaction, it is preferable to use a reactor in which the reaction heat can be easily controlled. Therefore, a fixed bed reactor or an adiabatic quenching reactor can be used. In addition, since coke is generated in the reaction and deactivation of the catalyst occurs due to the produced coke, a system capable of regenerating the catalyst by cycling the coke during the reaction is needed. Therefore, it is preferable to use a swing bed reactor.

The reaction of the third step, the reaction temperature of 250 ℃ to 500 ℃, a reaction pressure of 2bar to 30bar and a space velocity of 30,000h 1,000h ^-1 to ^- is preferably carried out in ^one of the conditions.

If the reaction temperature is lower than 250 ° C, there is a problem that the conversion of ethylene is lowered. When the temperature exceeds 500 ° C, paraffinic hydrocarbons and aromatics are increased, coke production is increased on the catalyst, . Further, when the reaction pressure is less than 2 bar, there is a problem that the conversion of ethylene is lowered. As the pressure is higher, the oligomerization reaction as the reaction is dominant and the conversion of ethylene is increased. When the reaction pressure is lower, cracking reaction of the heavy olefin And the parallel conversion rate drops. When the reaction pressure exceeds 30 bar, there is a problem in that a lot of operation cost is required for compressing the reactant.

The ethylene conversion in the third step may be 70% or more, preferably 80% or more, and more preferably 90% or more.

The composition discharged through the reactor of the third step may include CO, CO ₂ , H ₂ O, hydrogen, methane, C ₂ hydrocarbons, C ₃ to C ₄ hydrocarbons, C 5 + hydrocarbons, aromatics (aromatics) and the like.

The C ₂ hydrocarbons discharged through the reactor in the third step may include ethane and ethylene. In order to separate the ethane and ethylene from the downstream process, a complex process such as the multi-stage distillation column is required, It is necessary to increase the ethylene conversion rate in order to exclude it. Therefore, the present invention enables the recycling to the first step without separating them separately by identifying the optimum reaction conditions and catalysts capable of maximizing the conversion of ethylene in the third step reaction.

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 dimerization reaction of methane after the third step.

In the present invention, the product of the third step may be subjected to gas-liquid separation, and the gas phase of the gas-liquid separated second step product may be separated and recycled to the first step. At this time, unreacted methane is the main component in meteoric quantification of methane in the first stage. The recycled meteoric components to the first stage can be distributed to the oxidative quantification reaction of methane or the steam reforming reaction of methane, or alternatively can be distributed in a proportion to all of the reactions.

Methane conversion rate of about 30% and 100% conversion ratio are obtained when the methane to oxygen ratio is about 3: 1. If the C ₂ yield is about 15 to 20%, the reaction heat of the methanation oxidation reaction The methane required for the steam reforming reaction of methane, which is needed to neutralize the methane, is about half of the methane that is used to quantify the oxidation of methane. The conversion rate of the methane reforming reaction is about 95%. Therefore, as shown above, it is necessary to recycle part of the recycled methane to the quantification reaction of methane because the unreacted methane is excessively recycled as a whole by the reforming reaction of methane.

On the other hand, the by-products of the methanation oxidation reaction include CO, H ₂ , H ₂ O, and CO ₂ , which are products or reactants of the reforming reaction of methane and therefore are recycled to the methane reforming reaction It is advantageous. In conclusion, the amount and composition of by-products in the methanation of methane oxidation are advantageous in terms of the efficiency of the whole process, depending on the reaction conditions, but the recycle of the by-products, which are recycled first as a reforming reaction of methane, to the oxidation reaction of methane.

The product from the olefin oligomerization reactor in the third step may be cooled and separated into a gas phase and a liquid phase using a gas-liquid separator. The liquid product consists of C5 + hydrocarbons, including C5 + olefins, and water, the product of the oxidation-reduction reaction of methane. C5 + hydrocarbons can be obtained as gasoline by an additional hydrogenation process and can be converted to light olefins by cracking processes.

In the methane conversion method of the present invention, the conversion of methane is 20 to 80%, the selectivity to C2 + hydrocarbons is 40 to 80%, and the total yield of C2 + hydrocarbons is 15 to 25%.

In a second aspect of the present invention, there is provided an exhaust gas recirculation system comprising an exothermic reaction channel and an endothermic reaction channel, wherein a temperature (T ₁ ) of an exothermic reaction channel in which a quantitative oxidation reaction of methane is performed, The first step of performing the oxidation quantification reaction of methane in the exothermic reaction channel and the steam reforming reaction of methane in the endothermic reaction channel in the heat exchange reactor in which heat transfer from the exothermic reaction channel to the endothermic reaction channel is higher than the temperature (T ₂ ) ; A second step of converting the synthesis gas formed in the steam reforming reaction of methane in the first step; A third step of producing a liquid hydrocarbon product by an oligomerization reaction from C2 hydrocarbons formed in the dimerization reaction of methane in the first step; And a fourth step of recycling the unreacted gas in the third step to the steam reforming reaction and the oxidation reaction of methane in the first step.

The heat exchange reactor may be a microchannel reactor.

Each channel of the microchannel reactor may be filled with a catalyst for quantification of methane oxidation or a catalyst for steam reforming reaction of methane.

The exothermic reaction channel in which the oxidation quantification reaction of methane is performed and the endothermic reaction channel in which the steam reforming reaction of methane is performed may be disposed adjacent to each other.

Hereinafter, embodiments of the present invention will be described in detail with reference to the drawings as necessary. The following embodiments are illustrative of the present invention and are not intended to limit the present invention. FIG. 1 shows a schematic diagram of a process for producing liquid hydrocarbons or compounds in methane. The process according to the present invention consists largely of three reaction processes. First, the methane-containing gas is introduced into a microchannel reactor to convert methane, and oxygen is supplied as a reactant to produce a C ₂ compound (ethane + ethylene) by the oxidation reaction of methane as the first reaction Simultaneously, a synthesis gas is produced by a reforming reaction of methane as a second reaction; Introducing a gas containing ethylene as a product of the first reaction into an olefin oligomerization reactor to produce C5 + hydrocarbons; And a synthesis gas, which is a product of the second reaction, into a synthesis gas conversion reactor and converting the synthesis gas into synthetic oil, methanol or hydrogen, and further includes recovery and recycling processes of each process product.

2 is a process schematic diagram showing the conversion process of the methane-containing gas in more detail, which shows a process according to the present invention in which C5 + hydrocarbons and synthetic oil or methanol or hydrogen are produced in methane and the energy consumption and the operating cost are lower . First, the reaction gas containing methane flows into the oxidation-reduction reaction layer 2 of methane in the microchannel reactor 1 through the flow 21, and oxygen is also supplied as an oxidizing agent for methane. At this time, an unreacted recycled methane recycle stream 27 is mixed with the stream 21 and fed to the methane oxidation determination reaction layer 2. After the oxidation quantification reaction of methane, the product is discharged as stream 22 and quenched by cooling water, whereupon the cooling water is heated to additionally obtain high pressure steam. The cooled stream 22 flows into a flash column 4, which is a gas-liquid separator, to separate the phases, and the liquid water is discharged into the flow 24. The condensed and discharged water can be reused as process water after the purification process. The remaining gaseous gas component is stream 23, boosted or adjusted at a pressure of at least 3 atmospheres suitable for the ethylene oligomerization reaction, and then subjected to heat exchange to enter the olefin oligomerization reactor 5. After the olefin oligomerization reaction, the product is discharged to stream 25 and introduced into separator 6 to separate the phases. The gas phase separated in the separator 6 is recycled to the microchannel reactor 1 as stream 27.

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

The present invention relates to a process for the simultaneous reduction of the oxidation reaction of methane as an exothermic reaction and the steam reforming reaction of methane as an endothermic reaction to exchange heat of reaction to achieve a thermal neutral reaction condition to improve the thermal efficiency of the process, By providing a high efficiency total integrated process, it is possible to provide a method for more efficient conversion of methane.

Figure 1 is a schematic diagram of a process for producing liquid hydrocarbons or compounds in methane.
FIG. 2 is a process overview diagram showing the methane-containing gas conversion step in more detail.
FIG. 3 is a microchannel reactor in which the oxidation-reduction reaction of methane and the steam reforming reaction of methane occur; And an exothermic reaction and an endothermic reaction in a channel, and a channel heat exchange.

Hereinafter, the present invention will be described in detail with reference to the following examples. However, the present invention is not limited to these examples.

Manufacturing example  1: Microchannel Reactor Fabrication

Inconel 600 was used as a heat resistant material for the microchannel reactor. In the microchannel reactor, inconel plates are stacked, and each layer has a certain thickness and is separated from each other. The joining of the reactor is carried out by using diffusion bonding, brazing welding, welding (laser, arc and gas welding, etc.), which is a method of applying heat and pressure at a time after stacking the plates . The microchannel reactor is provided with a reaction layer in which the oxidation quantification reaction of methane and the reforming reaction of methane are alternately performed. The microchannel reactor is composed of three layers of oxidation quantification reaction layer, four layers of methane reforming reaction layer, and a plate size of 6 cm x 6 cm. Each layer is filled with respective catalysts, and a filter It is built-in. The thickness, shape and structure of the catalyst layer can be designed in various ways, and can be designed differently depending on the amount of endothermic reaction of methane oxidation and methane reforming. The methanation oxidation reaction layer and methane reforming reaction layer of the microchannel reactor were each equipped with a thermocouple to monitor the internal temperature of the reaction layer.

FIG. 3 is a schematic diagram showing a microchannel reactor in which the oxidation quantification reaction of methane and the steam reforming reaction of methane are performed as described above. In the microchannel reactor, the laminate for exothermic reaction and the laminate for endothermic reaction are alternately laminated so that the endothermic reaction laminate can receive heat supply from the laminate for exothermic reaction laminated adjacent to each other.

Manufacturing example  2: Methane Oxidation quantity  Preparation of Catalyst for Reaction

Catalysts for quantitative oxidation of methane were prepared using commercial silica (SiO ₂ , Davisil grade 635) as a catalyst carrier. First, La (NO ₃ ) ₃ .6H ₂ O, Mn (NO ₃ ) ₂ .4H ₂ O and Na ₂ WO ₄ .2H ₂ O were dissolved in distilled water as La, Mn and Na ₂ WO ₄ source The aqueous solution was supported on a silica carrier by incipient impregnation. The supported catalyst was dried and calcined at 800 ° C for 5 hours to be used as a catalyst for quantitative oxidation of methane. The composition of the prepared catalyst is as 4.75Na ₂ WO ₄ /2Mn/0.25La/SiO ₂ number represents the% by weight.

Manufacturing example  3: Methane For reforming reaction  Catalyst preparation

A catalyst for the reforming reaction of methane was prepared by using γ-alumina (γ-Al ₂ O ₃ ) as a catalyst carrier. First, an aqueous solution prepared by dissolving Tetraammineplatinum (II) hydroxide hydrate, Ni (NO ₃ ) ₂ · 6H ₂ O and Mg (NO ₃ ) ₂ · 6H ₂ O in distilled water as a source of Pt, Ni and Mg was initially wetted And then impregnated with alumina in an incipient impregnation method. The supported catalyst was dried and calcined at 900 ° C. for 5 hours to be used as a reforming catalyst for methane. The composition of the prepared catalyst is 0.05 Pt / 12Ni-5 Mg / Al ₂ O ₃ , and the numbers indicate weight%.

Manufacturing example  4: Ethylene For oligomerization  Catalyst preparation

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

Manufacturing example  5: Fisher- Tropsch  Preparation of cobalt-based catalyst for reaction

30.0 g of gamma-alumina was added to a solution of 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 12 hours or more and then calcined at 400 ° C. in an air atmosphere for 5 hours to prepare a 15 wt% Co / Al ₂ O ₃ catalyst.

Manufacturing example  6: Fisher- Tropsch  For reaction Iron-based  Catalyst preparation

In distilled water _{500ml, Fe (NO 3) 3} · 9H 2 O 100g, Cu (NO 3) 2 · 6H 2 O 2.39g, Al (NO 3) 3 · 6H 2 O 18.57g and Mn (NO _{3) 2} · 6H ₂ O was prepared to prepare an aqueous metal solution. An aqueous solution prepared by dissolving 66.36 g of K ₂ CO _{3 in} 300 ml of distilled water was prepared. The aqueous solution was placed in a beaker containing 1000 ml of distilled water heated to 70 ° C, stirred at the same time, . It was stirred at the same temperature for 3 hours and then cooled to room temperature (20 DEG C). The resulting precipitate was mixed with distilled water three times that of the precipitated slurry and precipitated again. The filtrate was discarded and the above procedure was repeated two more times and filtered. After filtration, the slurry was extruded by an extruder, and then dried at room temperature for 12 hours and at 110 DEG C for 12 hours, and then calcined at 400 DEG C for 6 hours. KNO a catalyst precursor of the results, the distilled water and 30ml ₃ , Dried at 110 ° C for 12 hours and then calcined at 400 ° C for 6 hours to obtain an iron-based Fischer-Tropsch catalyst having a composition of 2.5K / 100Fe / 4Cu / 10Mn / 20Al ₂ O ₃ molar ratio .

Manufacturing example  7: Preparation of Catalyst for Methanol Synthesis

A metal aqueous solution was prepared by dissolving 18.0 g of Cu (NO ₃ ) ₂ , 29.55 g of Zn (NO ₃ ) ₂ .6H ₂ O and 9.5 g of Al (NO ₃ ) ₃ .9H ₂ O in 300 ml of distilled water, K ₂ CO ₃ , And then they were placed in a beaker having 500 ml of distilled water heated to 70 ° C and mixed and co-co-treated with each other at the same time. It was stirred at the same temperature for 3 hours and then cooled to room temperature (20 DEG C). The resulting precipitate was mixed with distilled water three times that of the precipitated slurry and precipitated again. The filtrate was discarded and the above procedure was repeated two more times and filtered. After filtration, the slurry was extruded by an extruder, and then dried at room temperature for 12 hours and at 110 ° C for 12 hours, and then calcined at 300 ° C for 3 hours. The catalyst obtained above had a methanol synthesis catalyst having a composition of 0.75 Cu / 1 Zn / 0.26 Al oxide molar ratio.

Example  One

The catalyst for quantitative oxidation of methane and the reforming catalyst for methane prepared in Preparation Examples 2 and 3 were installed in the microchannel reactor manufactured in Production Example 1, and the exothermic reaction and the endothermic reaction were thermally neutralized by heat exchange Respectively. The volume of the catalyst mounted on the oxidation-reduction reaction layer of methane was 23 cc, and the volume mounted on the reforming reaction layer of methane was 28 cc. The microchannel reactor equipped with the catalyst was heated to an external temperature of 780 DEG C by an electric furnace. At this time, the inside of the reactor was allowed to flow at a flow rate of 300 cc / 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 700 cc / min of methane, 500 cc / min of nitrogen and 1.4 cc / min of water, and the reaction inlet pressure is 0.4 bar. The amount of gas in the oxidation quantification reaction of methane was 1500 cc / min of methane, 3000 cc / min of nitrogen (GC standard gas and dilution gas) and 600 cc / min of oxygen, and the reaction inlet pressure was 1.8 bar. The temperature inside the methanation oxidation reaction layer was 795 ° C, and the temperature inside the reforming reaction layer of methane was 723 ° C. It was higher than the outer layer temperature of the microchannel reactor due to the exothermic reaction of the oxidation reaction of methane. On the other hand, the temperature inside the reforming reaction layer of methane was lower than that of the microchannel reactor due to the endothermic reaction. The composition of the product was analyzed by gas chromatography (GC). Before and after the reaction, methane was detected by GC-TCD (thermal conductive detector) and methane conversion was calculated based on the internal standard gas, nitrogen. Hydrocarbon products were separated by GasPro column and detected by FID (Flame ionization detector). Product selectivity was calculated based on methane.

The results of the oxidation reaction of methane obtained above are shown in Table 1 below, and the results of the reforming reaction of methane are shown in Table 2 below.

Example  2

3Na ₂ WO ₄ /2Mn/0.5La/SiO ₂ catalyst was prepared by the same method as in Preparation Example 2, except that the methane reforming catalyst was the same catalyst as in Production Example 2 Were used. The catalyst was mounted in the microchannel reactor manufactured in Experimental Example 1, and the exothermic reaction and the endothermic reaction were thermally neutralized by heat exchange. The volume of the catalyst loaded on the oxidation-reduction reaction layer of methane was 20 cc, and the volume mounted on the reforming reaction layer of methane was 28 cc. The microchannel reactor equipped with the catalyst was heated to an external temperature of 780 DEG C by an electric furnace. At this time, the inside of the reactor was allowed to flow at a flow rate of 300 cc / min. When the outside temperature of the reactor is increased to 780 degrees, the gas composition and flow rate of the reforming reaction layer are gradually changed. The amount of gas in the normal reforming reaction is 700 cc / min of methane, 500 cc / min of nitrogen (GC standard gas) and 1.4 cc / min of water, and the reaction inlet pressure is 0.4 bar. The amount of gas in the methane oxidation quantification reaction is 1300 cc / min of methane, 2600 cc / min of nitrogen (GC standard gas and dilution gas) and 520 cc / min of oxygen, and the reaction inlet pressure is 1.8 bar. The temperature inside the oxidation-reduction reaction layer of methane was 818 ° C, and the temperature inside the reforming reaction layer of methane was 710 ° C. It was higher than the outer layer temperature of the microchannel reactor due to the exothermic reaction of the oxidation reaction of methane. On the other hand, the temperature inside the reforming reaction layer of methane was lower than that of the microchannel reactor due to the endothermic reaction.

The results of the oxidation reaction of methane obtained above are shown in Table 1 below, and the results of the reforming reaction of methane are shown in Table 2 below.

Example  3

3Na ₂ WO ₄ /2Mn/0.5La/SiO ₂ catalyst was prepared by the same method as in Preparation Example 2, except that the methane reforming catalyst was the same catalyst as in Production Example 2 Were used. The catalyst was mounted in the microchannel reactor manufactured in Production Example 1, and the exothermic reaction and the endothermic reaction were thermally neutralized by heat exchange. The volume of the catalyst loaded on the methanation oxidation reaction layer was 20 cc, and the volume mounted on the reforming reaction layer of methane was 28 cc. The microchannel reactor equipped with the catalyst was heated to an external temperature of 780 DEG C by an electric furnace. At this time, the inside of the reactor was allowed to flow at a flow rate of 300 cc / 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 800 cc / min of methane, 500 cc / min of nitrogen (GC standard gas) and 1.6 cc / min of water, and the reaction inlet pressure is 0.5 bar. The amount of the gas in the quantification reaction of methane was 1100 cc / min of methane, 1833 cc / min of nitrogen (GC standard gas and dilution gas), and 436 cc / min of oxygen, and the reaction inlet pressure was 1.1 bar. The temperature inside the oxidation-reduction reaction layer of methane was 782 ° C, and the temperature inside the reforming reaction layer of methane was 778 ° C. The temperature inside the reforming reaction layer of methane was lower than that of the microchannel reactor due to endothermic reaction due to the exothermic reaction of the oxidation reaction of methane.

The results of the oxidation reaction of methane obtained above are shown in Table 1 below, and the results of the reforming reaction of methane are shown in Table 2 below.

Comparative Example  One

3Na ₂ WO ₄ /2Mn/0.5La/SiO ₂ catalyst was prepared and used in the same manner as in Production Example 2, except that a catalyst for quantitative oxidation of methane was prepared. The catalyst was mounted in an inconel tube (outer diameter 1/2 inch) reactor to carry out the oxidation quantification reaction of methane. Unlike Examples 1 to 3, only the oxidation reaction of methane was performed using a single tube. The weight of the catalyst loaded in the dimerization reaction of methane was 0.2 g. The reactor equipped with the catalyst was heated to an external temperature of 800 DEG C by an electric furnace.

The amounts of the gases in the methane oxidation quantification reaction were 166 cc / min of methane, 139 cc / min of nitrogen (GC standard gas and dilution gas) and 55 cc / min of oxygen, and the reaction inlet pressure was 0.2 bar. The temperature inside the oxidation-quantification reaction layer of methane was 850 ° C, but the internal temperature rose due to the exothermic reaction of the oxidation reaction of methane oxidation.

The results of the oxidation quantification reaction of methane obtained above are shown in Table 1 below. The results are similar to those of the examples but show only a similar yield of C ₂ , but only the oxidation of methane. Methane reforming reaction at the same time, when the flow rate of the entire reacting gas is increased due to the increase in the amount of catalyst and the increase in the amount of the catalyst, the temperature of the reactor is increased by 900 ° C or more to lower the C ₂ yield, Can also cause serious hazards. On the other hand, when a microchannel reactor having a high heat exchange efficiency is disposed alternately in layers for performing the quantitative oxidation reaction of methane and the reforming reaction of methane as in the embodiment, the reaction heat can be easily controlled even if the reactor size is increased, There is an advantage that heat can be effectively used as a heat source of an endothermic reaction. In the practical example, the temperature of the reactor can be kept constant even if the amount of catalyst is increased 30 times or more as compared with the comparative example.

Comparative Example  2

5Na ₂ WO ₄ /2Mn/0.5La/SiO ₂ catalyst was prepared and used in the same manner as in Production Example 2, except that a catalyst for quantitative oxidation of methane was prepared. The catalyst was mounted in a quartz tube (outer diameter 3/4 inch) reactor to carry out the oxidation quantification reaction of methane. Unlike Examples 1 to 3, only the oxidation reaction of methane was performed using a single tube. The weight of the catalyst loaded in the dimerization reaction of methane was 0.5 g. The reactor equipped with the catalyst was heated to an external temperature of 800 DEG C by an electric furnace.

The amount of gas in the methane oxidation quantification reaction is 166 cc / min of methane, 139 cc / min of nitrogen (GC standard gas and dilution gas) and 55 cc / min of oxygen, and the reaction inlet pressure is 0.2 bar. The temperature inside the oxidation-quantification reaction layer of methane was 920 占 폚. The internal temperature greatly increased due to the exothermic reaction of the oxidation reaction of methane.

The results of the oxidation quantification reaction of methane obtained above are shown in Table 1 below. Unlike Comparative Example 1, when the amount of catalyst is increased from 0.2 g to 0.5 g and the material of the reactor is changed from a metal having a high thermal conductivity to a quartz reactor having a low thermal conductivity, the temperature inside the reactor is increased from 850 to 920 As shown in Fig. Therefore, if the amount of the catalyst is further increased, the temperature control of the reactor may be difficult and the C ₂ yield may be lowered. On the other hand, when the oxidation reaction of methane and the reforming reaction of methane are combined as in the present invention, the problem of temperature control and the low thermal efficiency of the process can be overcome.

Product distribution of oxidation quantification reaction of methane product Example 1 Example 2 Example 3 Comparative Example 1 Comparative Example 2 Methane conversion (%) 34.5 32.8 33.5 31.7 33.7 Oxygen conversion (%) 98.8 97 93.2 93.2 98.1 Product selection%) 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 distribution of reforming reaction of methane product Example 1 Example 2 Example 3 Methane conversion (%) 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 the catalyst for the olefin (ethylene) oligomerization reaction, the zeolite catalyst (HZSM-5, Si / Al = 25 mole ratio, 0.7 wt% P supported) prepared in Preparation Example 4 was used. The reaction was carried out in a fixed-bed reactor packed with a catalyst, and the reactant used in the reaction was a simulated gas having a composition similar to that of the oxidation-reduction reaction of methane. (Composition of simulated gas: 5.0% nitrogen, 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 4,000 h ^-1 to convert ethylene to C 5 + hydrocarbons. The composition of the product was analyzed by gas chromatography (GC). Before and after the reaction, ethylene was detected with a GC-TCD (thermal conductive detector) and ethylene conversion was calculated based on the internal standard gas, nitrogen. Hydrocarbon products were separated by GasPro column and detected by FID, and the product selectivity was calculated on the basis of ethylene. The product distribution of the ethylene oligomerization reaction is shown in Table 3 below.

Example  5

The same procedure as in Example 4 was carried out except that HZSM-5 (Si / Al = 50 mole ratio) catalyst was used and the reaction temperature was 380 ° C., the reaction pressure was 7 bar and the space velocity (GHSV) was 4,000 h ^-1 (Ethylene) oligomerization reaction to convert ethylene to C5 + olefin. The results are shown in Table 3 below.

Product distribution of ethylene oligomerization reaction 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 conducted to produce synthetic oil by Fischer-Tropsch reaction in syngas. A 0.5-inch stainless steel fixed-bed reactor was charged with 0.5 g of the cobalt-based catalyst prepared in Preparation Example 5 and subjected to reduction treatment in an atmosphere of hydrogen (5 vol% H ₂ / He) at 400 ° C for 5 hours. Reaction temperature 230 ℃, a reaction pressure of 20 kg / cm ^2, The molar ratio of carbon monoxide: hydrogen: argon (internal standard) was fixed at a rate of 31.5: 63.0: 5.5 at a space velocity of 4000 L / kg _cat / hr and injected into the reactor. The Fischer-Tropsch reaction was carried out and the activity of the catalyst after 40 hours of reaction was measured. The results are shown in Table 4 below.

Example  7

The iron-based Fischer-Tropsch catalyst prepared in Preparation Example 6 was pulverized and classified to a size of about 1 mm, and 1 g of the thus-obtained iron-based Fischer-Tropsch catalyst was charged into a one-stage fixed bed reactor and activated at a normal pressure and 450 DEG C for 12 hours in a hydrogen atmosphere . Reaction temperature 280 ℃, a reaction pressure of 10 kg / cm ^2, The molar ratio of carbon monoxide: hydrogen: carbon dioxide: argon (internal standard) was fixed at a rate of 18.0: 60.5: 16.0: 5.5 at a space velocity of 3600 L / kg _cat / hr. The Fischer-Tropsch reaction was carried out and the activity of the catalyst after 40 hours of reaction was measured. The results are shown in Table 4 below.

Example  8

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

Product distribution of Fischer-Tropsch reaction division catalyst CO conversion (%) Hydrocarbon selectivity (%) CH ₄ C ₂ -C ₄ C ₅ + Example 6 15% Co /? - 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

1: Microchannel reactor
2: Oxidation of Methane in Microchannel Reactor
3: Methane reforming reaction layer in a microchannel reactor
4: gas-liquid separator
5: Olefin oligomerization reactor
6, 8: Separation device
7: Synthetic gas conversion reactor
21: Piping for supplying methane and oxygen to the oxidation quantification reaction of methane
22: piping for supplying the oxidation-reduction reaction product of methane to the gas-liquid separator
23: piping for supplying the gas component separated in the gas-liquid separator to the olefin oligomerization reactor
24: piping for recovering the liquid component separated from the gas-liquid separator
25: piping for supplying the olefin oligomerization reaction product to the separation apparatus
26: piping for recovering the liquid component from the separation device
27: Piping that separates gas components from the separator and recirculates them to the microchannel reactor
31: Piping for supplying methane and water vapor to the reforming reaction of methane
32: piping for supplying the reforming reaction product of methane to the syngas converter
33: piping for supplying the product of the syngas converter to the separator
34: Piping for separating unreacted components from the separation unit and recirculating it to the synthesis gas converter
35: piping for recovering the product from the separation apparatus

Claims (14)

An exothermic reaction channel and an endothermic reaction channel are provided so that the temperature T ₁ of the exothermic reaction channel in which the oxidation reaction of methane is being performed is higher than the temperature T ₂ of the endothermic reaction channel in which the steam reforming reaction of methane is performed And a first step of performing the oxidation quantification reaction of methane in the exothermic reaction channel and performing the steam reforming reaction of methane in the endothermic reaction channel in the heat exchange reactor in which heat is transferred from the exothermic reaction channel to the endothermic reaction channel, . The methane conversion method according to claim 1, wherein the methane-containing gas as a reactant to be injected into the heat exchange reactor in the first stage is natural gas. The method of claim 1, wherein the synthesis gas is formed as a product by the steam reforming reaction of methane in the first step. The method according to claim 1, further comprising a second step of producing synthetic oil from the synthesis gas formed in the first step by methanol synthesis, hydrogen production, ammonia production, or Fischer-Tropsch reaction, . The method of claim 1, wherein ethylene and / or ethane is formed as the C2 hydrocarbon as the product by the oxidation-reduction reaction of methane in the first step. The method of claim 5, further comprising a third step of producing a liquid hydrocarbon product by an ethylene oligomerization reaction from the C2 hydrocarbon formed in the first step. 7. The method of claim 6, further comprising a fourth step of recycling the unreacted gas of the third step to the steam reforming reaction of the first step and the dimerization reaction of methane. The methane conversion method according to claim 1, wherein the exothermic reaction channel in which the oxidation reaction of methane is performed and the endothermic reaction channels in which the steam reforming reaction of methane is performed are alternately disposed adjacent to each other. The methane conversion method according to any one of claims 1 to 8, wherein the heat exchange reactor in the first stage is a microchannel reactor. The method of claim 9, wherein each channel of the microchannel reactor is filled with a catalyst for quantification of methane oxidation or a catalyst for steam reforming reaction of methane. An exothermic reaction channel and an endothermic reaction channel are provided so that the temperature T ₁ of the exothermic reaction channel in which the oxidation reaction of methane is being performed is higher than the temperature T ₂ of the endothermic reaction channel in which the steam reforming reaction of methane is performed A first step of performing the oxidation quantification reaction of methane in an exothermic reaction channel and the steam reforming reaction of methane in an endothermic reaction channel in a heat exchange reactor in which heat is transferred from an exothermic reaction channel to an endothermic reaction channel;
A second step of converting the synthesis gas formed in the steam reforming reaction of methane in the first step;
A third step of producing a liquid hydrocarbon product by an oligomerization reaction from C2 hydrocarbons formed in the dimerization reaction of methane in the first step; And
And a fourth step of recycling the unreacted gas in the third step to the steam reforming reaction in the first step and the dimerization reaction in methane. 12. The method of claim 11, wherein the first stage heat exchange reactor is a microchannel reactor. 13. The method of claim 12, wherein each channel of the microchannel reactor is filled with a catalyst for quantification of methane oxidation or a catalyst for steam reforming reaction of methane. The methane conversion method according to claim 11, wherein the exothermic reaction channel in which the oxidation reaction of methane is performed and the endothermic reaction channels in which the steam reforming reaction of methane is performed are alternately disposed adjacent to each other.

Images