Method for producing butene and light aromatic hydrocarbon from dry gas

Abstract

The invention relates to a method for producing butylene and light aromatic hydrocarbon from dry gas, which comprises the following steps: introducing a dry gas raw material into a first superposition reactor to contact with a first superposition catalyst to carry out first reaction, and carrying out first separation on a first reaction product obtained by the reaction to obtain dry gas, butylene and a first liquid product; the first reaction comprises a dimerization reaction; introducing the dry gas obtained by the first separation into a second superposition reactor, contacting the dry gas with a second superposition catalyst to carry out a second reaction, and carrying out second separation on a second reaction product obtained by the reaction to obtain dry gas, liquefied gas, gasoline and a second liquid product; the second reaction comprises an oligomerization reaction, a cyclization dehydrogenation reaction and an aromatization reaction. According to the invention, the conversion rate of ethylene and the yields of butene and light aromatic hydrocarbons are improved, the proportion of n-butene in the obtained butene product is high, the utilization way of ethylene is widened, and high-value utilization of ethylene is realized.

Description

Technical Field

[0001] This invention belongs to the field of petrochemicals, specifically relating to a method for producing butene and light aromatics from dry gas. Background Technology

[0002] Dry gas, a byproduct of catalytic cracking, contains abundant ethylene; for example, the ethylene mass fraction in catalytic cracking dry gas ranges from 20% to 55%. However, due to insufficient attention paid to the ethylene in dry gas, it is usually treated as fuel gas. Therefore, how to utilize the ethylene in dry gas in a high-value manner has attracted much attention.

[0003] CN106278795A discloses a method for preparing ethylbenzene by alkylation of ethylene and benzene in dry gas, but this technology is constrained by the supply of benzene as a raw material. CN105985801A discloses a method for promoting the production of gasoline and liquefied petroleum gas (LPG) from catalytic dry gas. This method uses a mixture of rare earth-ZSM-5 / ZSM-11 co-crystallized zeolite and alumina to prepare a catalyst, which is then used for aromatization of dry gas and steam as co-feeds to produce gasoline and LPG. The LPG yield can reach 30%, and the gasoline liquid yield is approximately 50%.

[0004] CN102851063A discloses a method for producing high-octane clean gasoline through aromatization of dry gas and liquefied petroleum gas (LPG). The method involves loading an LPG aromatization catalyst at the top of a fixed-bed reactor and a dry gas aromatization catalyst at the bottom. The LPG feedstock, containing olefins, enters the reactor from the top and reacts with the LPG aromatization catalyst to undergo olefin chelation, cyclization, and aromatization reactions to generate aromatics. Meanwhile, cold dry gas enters the reactor from between the packing layers and reacts with the dry gas aromatization catalyst to generate aromatic oil. The butene yield obtained using this method is approximately 6%, and the C20 is... 5+ The yield of gasoline was 38.55%.

[0005] However, in existing technologies for producing liquefied petroleum gas and light aromatics from ethylene in dry gas, the yield of butene and the selectivity of n-butene need to be improved. Summary of the Invention

[0006] The purpose of this invention is to solve the problem of low butene yield and to further improve the selectivity of n-butene.

[0007] To achieve the above objectives, the present invention provides a method for producing butene and light aromatics from dry gas, the method comprising: The dry gas feedstock is introduced into the first superimposed reactor and comes into contact with the first superimposed catalyst to carry out the first reaction. The first reaction product obtained from the reaction is subjected to the first separation to obtain dry gas and butene. The first reaction includes a dimerization reaction. The dry gas obtained from the first separation is introduced into the second superimposed reactor and contacted with the second superimposed catalyst to carry out a second reaction. The second reaction product obtained from the reaction is then subjected to a second separation to obtain dry gas, liquefied gas and gasoline. The second reaction includes oligomerization, cyclization dehydrogenation and aromatization reactions.

[0008] Optionally, the conditions for the first reaction include: a reaction temperature of 100-400 °C, preferably 150-300 °C; a reaction pressure of 0.1-5 MPa, preferably 0.5-3 MPa; and a mass hourly space velocity of 0.1-100 h⁻¹. -1 Preferably 0.3-30 h -1 The conditions for the second reaction include: a reaction temperature of 150-500 °C, preferably 200-450 °C; a reaction pressure of 0.1-5 MPa, preferably 0.5-4 MPa; and a mass hourly space velocity of 0.1-20 h⁻¹. -1 Preferably 0.3-5 h -1 .

[0009] Optionally, the first and second composite catalysts each independently comprise 15-70% by weight of a binder, 10-65% by weight of a matrix, and 20-75% by weight of an active component; wherein the active component is selected from one or more of amorphous silica-alumina, zeolite molecular sieves, and mesoporous materials, preferably zeolite molecular sieves; the matrix is ​​selected from one or more of kaolin, montmorillonite, diatomaceous earth, attapulgite, sepiolite, halloysite, hydrotalcite, bentonite, and rettore; and the binder is selected from one or more of silica, alumina, zirconium oxide, titanium oxide, and amorphous silica-alumina.

[0010] Optionally, the zeolite molecular sieve includes FAU-structured molecular sieves and / or MFI-structured molecular sieves; preferably, the zeolite molecular sieve includes FAU-structured molecular sieves and MFI-structured molecular sieves, and based on the total weight of the zeolite molecular sieves, the zeolite molecular sieves include 0-80% by weight of FAU-structured molecular sieves and 20-100% by weight of MFI-structured molecular sieves; the FAU-structured molecular sieves are selected from at least one of Y-type molecular sieves and modified Y-type molecular sieves; the MFI-structured molecular sieves are selected from ZSM-5 molecular sieves and / or ZRP molecular sieves.

[0011] Optionally, the zeolite molecular sieve is a FAU structure molecular sieve and / or an MFI structure molecular sieve modified with non-metallic elements and / or transition metal elements; the non-metallic elements are selected from phosphorus and / or boron; the transition metal elements are selected from at least one of iron, cobalt and nickel.

[0012] Optionally, the first composite reactor and the second composite reactor are each independently selected from one of a fixed-bed reactor, a stirred tank reactor, and a tower reactor.

[0013] Optionally, the ethylene content in the dry gas feedstock is 5-80% by weight, preferably 10-70% by weight; preferably, the dry gas feedstock comes from one or more of a catalytic cracking unit, a delayed coking unit, a hydrocracking unit, an ethylene plant steam cracking unit, and a chemical plant MTO unit.

[0014] Optionally, the method further includes: contacting the gasoline with hydrogen and a hydrogenation catalyst under hydrogenation conditions to perform hydrogenation refining to obtain refined gasoline; and performing solvent extraction on the refined gasoline to obtain benzene, toluene and xylene.

[0015] Optionally, the hydrorefining conditions include: a hydrogen partial pressure of 2.0-15.0 MPa; a temperature of 200-400 °C; a hydrogen-to-oil volume ratio of 100-1000; and a volume hourly space velocity of 2.0-5 h⁻¹. -1 The hydrogenation catalyst comprises a support and a hydrogenation active component supported on the support; the support is alumina and / or amorphous silica-alumina, and the hydrogenation active component includes a Group VIB metal and / or a Group VIII metal.

[0016] Optionally, the extraction conditions include: a temperature of 80-120 °C; a volume ratio of extraction solvent to refined gasoline of 2-6; and the extraction solvent is selected from one or more of sulfolane, N-methylpyrrolidone, diethylene glycol ether, triethylene glycol ether, tetraethylene glycol, and dimethyl sulfoxide.

[0017] Through the above technical solution, this invention synthesizes butene and light aromatics by performing a two-step catalytic superposition reaction on ethylene in dry gas. This allows for separate control of the superposition reaction conditions, improving the conversion rate of ethylene and the yield of butene and light aromatics. Furthermore, the butene product has a high proportion of n-butene, broadening the utilization pathways of ethylene and realizing the high-value utilization of ethylene.

[0018] Other features and advantages of the present invention will be described in detail in the following detailed description section. Attached Figure Description

[0019] The accompanying drawings are provided to further illustrate the invention and form part of the specification. They are used together with the following detailed description to explain the invention, but do not constitute a limitation thereof. In the drawings: Figure 1 This is a flowchart illustrating a method in one specific embodiment of the present invention.

[0020] Explanation of reference numerals in the attached figures: 1. First superimposed reactor; 2. First separation device; 3. Second superimposed reactor; 4. Second separation device; 5. Pipeline; 6. Pipeline; 7. Pipeline; 8. Pipeline; 9. Pipeline; 10. Pipeline; 11. Pipeline; 12. Pipeline; 13. Pipeline; 14. Pipeline. Detailed Implementation

[0021] The following provides a detailed description of specific embodiments of the present invention. It should be understood that the specific embodiments described herein are for illustrative and explanatory purposes only and are not intended to limit the scope of the invention.

[0022] This invention provides a method for producing butene and light aromatics from dry gas. The method includes: introducing dry gas feedstock into a first composite reactor and contacting it with a first composite catalyst to carry out a first reaction; performing a first separation on the first reaction product obtained from the reaction to obtain dry gas and butene; the first reaction includes a dimerization reaction. The dry gas obtained from the first separation is introduced into the second superimposed reactor and contacted with the second superimposed catalyst to carry out a second reaction. The second reaction product obtained from the reaction is then subjected to a second separation to obtain dry gas, liquefied gas and gasoline. The second reaction includes oligomerization, cyclization dehydrogenation and aromatization reactions.

[0023] This invention synthesizes butene and light aromatics through a two-step catalytic chelation reaction of ethylene in dry gas. This allows for separate control of the chelation reaction conditions, improving the conversion rate of ethylene and the yields of butene and light aromatics. Furthermore, the butene product has a high proportion of n-butene, broadening the utilization pathways of ethylene and maximizing the utilization of dry gas with relatively low added value, thus possessing broad application prospects.

[0024] In this invention, the dry gas feedstock refers to catalytic cracking dry gas, deep catalytic cracking dry gas, coking dry gas, and other ethylene-containing gases produced during secondary petroleum processing or coal processing. It is mainly composed of hydrogen, methane, ethane, ethylene, and propylene, and the ethylene content varies depending on the equipment and operating conditions.

[0025] In some embodiments of the present invention, the dry gas feedstock may come from one or more of a catalytic cracking unit, a delayed coking unit, a hydrocracking unit, an ethylene plant steam cracking unit, and a chemical plant MTO unit. Preferably, the dry gas feedstock is dry gas produced as a byproduct of a catalytic cracking unit and / or a delayed coking unit.

[0026] In some embodiments of the present invention, the ethylene content in the dry gas feedstock is 5-80% by weight, preferably 10-70% by weight.

[0027] The inventors of this invention discovered that allowing ethylene to undergo dimerization at a lower conversion rate can improve the selectivity of butene, while under reaction conditions that increase the ethylene conversion rate, ethylene can generate more aromatics. By controlling the reaction conditions of the first and second reactions respectively, butene and aromatic-rich gasoline can be obtained in high yields, respectively.

[0028] In some embodiments of the present invention, the conditions for the first reaction include: a reaction temperature of 100-400 °C; a reaction pressure of 0.1-5 MPa; and a mass hourly space velocity of 0.1-100 h⁻¹. -1 For example, the reaction temperature of the first reaction can be 100℃, 120℃, 140℃, 160℃, 180℃, 200℃, 210℃, 220℃, 230℃, 240℃, 250℃, 260℃, 270℃, 280℃, 290℃, 300℃, 310℃, 330℃, 350℃, 380℃, 400℃, or any temperature within the aforementioned range; the reaction pressure can be 0.1 MPa, 1 MPa, 1.5 MPa, 2.0 MPa, 3.0 MPa, 4.0 MPa, 5.0 MPa, or any pressure within the aforementioned range; the mass hourly space velocity can be 0.1 h⁻¹. -1 1.0 h -1 3.5 h -1 5.0 h -1 10.0 h -1 20 h -1 30 h -1 50 h -1 80 h -1 100 h -1 Or any mass hourly space velocity within the aforementioned range. Under the above reaction conditions, ethylene undergoes a dimerization reaction in the first reactor in contact with the first composite catalyst, with butene being the main product. Simultaneously, by controlling an appropriate reaction depth, the distribution of butene isomers can be regulated to maximize the proportion of n-butene in the total butene product. The products of the first reaction of ethylene also include a first liquid product with a distillation range of 45-350 °C.

[0029] In some preferred embodiments of the present invention, in order to further improve the yield of butene and the proportion of n-butene in the butene product, the conditions for the first reaction include: a reaction temperature of 150-300 °C; a reaction pressure of 0.5-3 MPa; and a mass hourly space velocity of 0.3-30 h⁻¹. -1 .

[0030] In some embodiments of the present invention, the conditions for the second reaction include: a reaction temperature of 150-500 °C; a reaction pressure of 0.1-5 MPa; and a mass hourly space velocity of 0.1-20 h⁻¹. -1 For example, the second reaction can be carried out at temperatures of 150℃, 180℃, 200℃, 210℃, 220℃, 230℃, 240℃, 250℃, 260℃, 270℃, 280℃, 290℃, 300℃, 310℃, 320℃, 330℃, 340℃, 350℃, 360℃, 370℃, 380℃, 390℃, 400℃, 420℃, 440℃, 450℃, 460℃, 480℃, 500℃, or any temperature within the aforementioned range; the reaction pressure can be 0.1 MPa, 1 MPa, 1.5 MPa, 2.0 MPa, 2.2 MPa, 3.0 MPa, 4.0 MPa, 5.0 MPa. Pressure of MPa or any pressure within the aforementioned range; mass hourly space velocity can be 0.1 h⁻¹. -1 1.0 h -1 1.1 h -1 1.3 h -1 3.5 h -1 5.0 h -1 10.0 h -1 15 h -1 20 h -1 Or any mass hourly space velocity within the aforementioned range. Under the above reaction conditions, ethylene in the second reactor mainly undergoes oligomerization, cyclization dehydrogenation, and aromatization reactions in contact with the second superimposed catalyst. Increasing the severity of the reaction is beneficial to increasing the yield of gasoline and the content of light aromatics in the gasoline. The products of the second reaction of ethylene also include a second liquid product, which is a fraction with a distillation range of 220-350 °C.

[0031] In some preferred embodiments of the present invention, in order to further improve the conversion rate of ethylene and the yield of light aromatics, the conditions for the second reaction include: a reaction temperature of 200-450 °C; a reaction pressure of 0.5-4 MPa; and a mass hourly space velocity of 0.3-5 h⁻¹. -1 .

[0032] In this invention, the method for separating dry gas, butene, gasoline and liquid products from the reaction products can employ conventional techniques in the art, and this invention does not limit this, so it will not be described in detail here.

[0033] In some embodiments of the present invention, the first composite catalyst and the second composite catalyst may be the same or different, and each independently comprises 15-70% by weight of binder, 10-65% by weight of matrix and 20-75% by weight of active component.

[0034] In some embodiments of the present invention, the active component is selected from one or more of amorphous silica-alumina, zeolite molecular sieves and mesoporous materials, preferably zeolite molecular sieves.

[0035] In some embodiments of the present invention, the zeolite molecular sieve includes FAU-structured molecular sieves and / or MFI-structured molecular sieves. Preferably, to improve the yield of butene and light aromatics, the zeolite molecular sieve may include both FAU-structured and MFI-structured molecular sieves. Specifically, based on the total weight of the zeolite molecular sieve, the zeolite molecular sieve may include 0-80% by weight of FAU-structured molecular sieves and 20-100% by weight of MFI-structured molecular sieves.

[0036] In some embodiments of the present invention, the FAU structure molecule is screened from at least one of Y-type molecular sieve and modified Y-type molecular sieve; the MFI structure molecule is screened from ZSM-5 molecular sieve and / or ZRP molecular sieve.

[0037] In some preferred embodiments of the present invention, the zeolite molecular sieve is a FAU structure molecular sieve and / or an MFI structure molecular sieve modified with non-metallic elements and / or transition metal elements. The non-metallic elements are selected from phosphorus and / or boron; the transition metal elements are selected from at least one of iron, cobalt, and nickel.

[0038] In some embodiments of the present invention, the matrix is ​​selected from one or more of kaolin, montmorillonite, diatomite, attapulgite, sepiolite, halloysite, hydrotalcite, bentonite, and rettoite.

[0039] In some embodiments of the present invention, the binder is selected from one or more of silicon oxide, aluminum oxide, zirconium oxide, titanium oxide, and amorphous aluminum silicate.

[0040] In some embodiments of the present invention, the first composite reactor and the second composite reactor are each independently selected from one of a fixed-bed reactor, a stirred tank reactor, and a tower reactor, preferably a fixed-bed reactor.

[0041] In some embodiments of the present invention, the superimposed reactor may be operated intermittently or continuously, preferably continuously.

[0042] In some embodiments of the present invention, the method further includes: hydrorefining the gasoline by contacting it with hydrogen and a hydrorefining catalyst under hydrogenation conditions to obtain refined gasoline; and solvent extraction of the refined gasoline to obtain benzene, toluene and xylene.

[0043] In some embodiments of the present invention, the conditions for hydrorefining may include: a hydrogen partial pressure of 2.0-15.0 MPa; a temperature of 200-400 °C; a hydrogen-to-oil volume ratio of 100-1000; and a volume hourly space velocity of 2.0-5 h⁻¹. -1 .

[0044] In some embodiments of the present invention, the hydrogenation catalyst comprises a support and a hydrogenation active component supported on the support; the support is alumina and / or amorphous silica-alumina, and the hydrogenation active component includes a Group VIB metal and / or a Group VIII metal. The method for hydrorefining gasoline can employ conventional techniques in the art, and this disclosure does not limit this process.

[0045] In some embodiments of the present invention, solvent extraction can be performed using conventional methods in the art, such as contacting refined gasoline obtained by hydrorefining with a solvent to extract aromatics; the extraction conditions include: a temperature of 80-120 °C; and a volume ratio of extraction solvent to refined gasoline of 2-6.

[0046] In some embodiments of the present invention, the extraction solvent is selected from one or more of sulfolane, N-methylpyrrolidone, diethylene glycol ether, triethylene glycol ether, tetraethylene glycol, and dimethyl sulfoxide. The solvent obtained from the solvent extraction separation can be recycled.

[0047] Figure 1 A flowchart illustrating a preferred embodiment of the present invention is shown below. Figure 1 The flowchart shown illustrates the method for producing butene and light aromatics from dry gas according to the present invention.

[0048] like Figure 1 As shown, ethylene-containing dry gas feedstock enters the first superposition reactor 1 via pipeline 5. Under the action of the first superposition catalyst, a dimerization reaction mainly occurs, and the resulting first reaction product enters the first separation unit 2 via pipeline 6 for separation, yielding dry gas, butene, and a first liquid product. Butene and the first liquid product are led out via pipelines 8 and 9, respectively, while unreacted dry gas is introduced into the second superposition reactor via pipeline 7. Under the action of the second superposition catalyst, oligomerization, cyclization dehydrogenation, and aromatization reactions occur, and the resulting second reaction product is introduced into the second separation unit 4 via pipeline 10. The second reaction product undergoes separation processes such as distillation and absorption in the second separation unit 4 to obtain dry gas, liquefied petroleum gas (LPG), a second liquid product, and superposition gasoline. Dry gas is led out via pipeline 11, LPG via pipeline 12, the second liquid product via pipeline 14, and superposition gasoline rich in light aromatics is led out via pipeline 13. The superposition gasoline undergoes subsequent hydrorefining and aromatics extraction to obtain light aromatics.

[0049] The present invention will be further described in detail below through examples, but the invention is not limited thereto. All raw materials used in the examples are commercially available.

[0050] In the embodiments of the present invention, both the first and second superimposed reactors are fixed-bed reactors. The ethylene content in raw material A used in the examples and comparative examples is 39.7% by volume, with the remainder being argon, used to simulate the dry gas composition of an industrial catalytic cracking unit.

[0051] The first and second superimposed catalysts in the embodiments and comparative examples of this invention are the same, produced by Sinopec Catalyst Co., Ltd., and the active components are Y-type molecular sieve and shape-selective ZSM-5 molecular sieve. Based on the total weight of the catalyst, the catalyst contains 10% by weight of Y-type molecular sieve and 25% by weight of shape-selective ZSM-5 molecular sieve. The properties of the catalyst are listed in Table 1.

[0052] Table 1

[0053] Example 1 according to Figure 1 The flowchart shows that ethylene-containing dry gas feedstock A enters the first superimposed reactor and undergoes a first reaction (mainly a dimerization reaction) under the action of a first superimposed catalyst. The first reaction product is then sent to a first separation unit for separation, yielding dry gas, butene, and a first liquid product. The conditions for the first reaction include: a reaction temperature of 250 °C, a reaction pressure of 1.5 MPa, and a mass hourly space velocity of 3.5 h⁻¹. -1 .

[0054] Unreacted dry gas from the first superimposed reactor is introduced into the second superimposed reactor, where a second reaction occurs under the action of the second superimposed catalyst. This second reaction includes oligomerization, cyclization dehydrogenation, and aromatization. The resulting products are separated to obtain dry gas, liquefied petroleum gas (LPG), gasoline, and a second liquid product. The conditions for the second reaction are: a reaction temperature of 350 °C, a reaction pressure of 2.0 MPa, and a mass hourly space velocity (HHSV) of 1.3 h⁻¹. -1 .

[0055] Gasoline was hydrorefined and aromatics extracted to obtain light aromatics. The hydrorefining conditions included a hydrogen partial pressure of 2.6 MPa, a temperature of 240 °C, a hydrogen-to-oil volume ratio of 400, and a volume hourly space velocity of 2.3 h⁻¹. -1 The conditions for aromatic hydrocarbon extraction include: a temperature of 90 °C, a volume ratio of extraction solvent to refined gasoline of 3.5, and the extraction solvent being sulfolane.

[0056] The product distribution is shown in Table 2. The distillation range of the liquid products is 220-350 °C.

[0057] Comparative Example 1 The comparative method for producing butene and light aromatics using dry gas is basically similar to that of Example 1, except that the dry gas feedstock A is directly introduced into the second superimposed reactor for the second reaction. The reaction conditions and product distribution are shown in Table 2.

[0058] Example 2 The method for producing butene and light aromatics using dry gas in this embodiment is basically similar to that in Example 1, except that the reaction temperature of the first reaction is 310 °C. The first reaction conditions and product distribution are shown in Table 2.

[0059] Example 3 The method for producing butene and light aromatics using dry gas in this embodiment is basically similar to that in Example 1, except that the reaction temperature of the second reaction is 460 °C. The reaction conditions and product distribution of the second reaction are shown in Table 2.

[0060] Example 4 Example 4 describes a method for producing butene and light aromatics using dry gas that is basically similar to that of Example 1, except that the reaction temperature for the first reaction is 200 °C and the reaction pressure is 2.0 MPa; the reaction temperature for the second reaction is 280 °C, the reaction pressure is 2.2 MPa, and the mass hourly space velocity (HHSV) is 1.1 h⁻¹. -1 The reaction conditions and product distribution are shown in Table 2.

[0061] Table 2

[0062] As can be seen from the results in the table above, the method provided by the present invention improves the conversion rate of ethylene in dry gas and significantly improves the yield of butene and the selectivity of n-butene.

[0063] The preferred embodiments of the present invention have been described in detail above. However, the present invention is not limited to the specific details in the above embodiments. Within the scope of the technical concept of the present invention, various simple modifications can be made to the technical solution of the present invention, and these simple modifications all fall within the protection scope of the present invention.

[0064] It should also be noted that the various specific technical features described in the above embodiments can be combined in any suitable manner without contradiction. To avoid unnecessary repetition, the present invention will not describe the various possible combinations separately.

[0065] Furthermore, various different embodiments of the present invention can be combined in any way, as long as they do not violate the spirit of the present invention, they should also be regarded as the content disclosed by the present invention.

Claims

1. A method for producing butene and light aromatics from dry gas, characterized in that, The method includes: The dry gas feedstock is introduced into the first superimposed reactor and comes into contact with the first superimposed catalyst to carry out the first reaction. The first reaction product obtained from the reaction is subjected to the first separation to obtain dry gas and butene. The first reaction includes a dimerization reaction. The dry gas obtained from the first separation is introduced into the second superimposed reactor and contacted with the second superimposed catalyst to carry out a second reaction. The second reaction product obtained from the reaction is then subjected to a second separation to obtain dry gas, liquefied gas and gasoline. The second reaction includes oligomerization, cyclization dehydrogenation and aromatization reactions.

2. The method according to claim 1, wherein, The conditions for the first reaction include: a reaction temperature of 100-400℃, preferably 150-300℃; a reaction pressure of 0.1-5 MPa, preferably 0.5-3 MPa; and a mass hourly space velocity of 0.1-100 h⁻¹. -1 Preferably 0.3-30 h -1 ; The conditions for the second reaction include: a reaction temperature of 150-500 °C, preferably 200-450 °C; a reaction pressure of 0.1-5 MPa, preferably 0.5-4 MPa; and a mass hourly space velocity of 0.1-20 h⁻¹. -1 Preferably 0.3-5 h -1 .

3. The method according to claim 1, wherein, The first and second composite catalysts each independently comprise 15-70% by weight of binder, 10-65% by weight of matrix and 20-75% by weight of active component; The active component is selected from one or more of amorphous silica-alumina, zeolite molecular sieves and mesoporous materials, preferably zeolite molecular sieves; The matrix is ​​selected from one or more of the following: kaolin, montmorillonite, diatomite, attapulgite, sepiolite, halloysite, hydrotalcite, bentonite, and rettoite. The binder is selected from one or more of silicon oxide, aluminum oxide, zirconium oxide, titanium oxide, and amorphous aluminum silicate.

4. The method according to claim 3, wherein, The zeolite molecular sieve includes FAU structure molecular sieve and / or MFI structure molecular sieve; Preferably, the zeolite molecular sieve includes FAU structure molecular sieve and MFI structure molecular sieve, and based on the total weight of the zeolite molecular sieve, the zeolite molecular sieve includes 0-80% by weight of FAU structure molecular sieve and 20-100% by weight of MFI structure molecular sieve. The FAU structure molecule is selected from at least one of Y-type molecular sieves and modified Y-type molecular sieves; The MFI structure molecules are screened from ZSM-5 molecular sieve and / or ZRP molecular sieve.

5. The method according to claim 3, wherein, The zeolite molecular sieve is a FAU structure molecular sieve and / or an MFI structure molecular sieve modified with non-metallic elements and / or transition metal elements. The non-metallic element is selected from phosphorus and / or boron; The transition metal element is selected from at least one of iron, cobalt, and nickel.

6. The method according to claim 1, wherein, The first and second composite reactors are each independently selected from one of a fixed-bed reactor, a stirred tank reactor, and a tower reactor.

7. The method according to claim 1, wherein, The ethylene content in the dry gas feedstock is 5-80% by weight, preferably 10-70% by weight. Preferably, the dry gas feedstock comes from one or more of the following: a catalytic cracking unit, a delayed coking unit, a hydrocracking unit, an ethylene plant steam cracking unit, and a chemical plant MTO unit.

8. The method according to claim 1, wherein, The method further includes: under hydrogenation conditions, contacting the gasoline with hydrogen and a hydrogenation catalyst to perform hydrogenation refining to obtain refined gasoline; and The refined gasoline was subjected to solvent extraction to obtain benzene, toluene, and xylene.

9. The method according to claim 8, wherein, The conditions for hydrorefining include: hydrogen partial pressure of 2.0-15.0 MPa; temperature of 200-400 °C; hydrogen-to-oil volume ratio of 100-1000; and volume hourly space velocity of 2.0-5 h⁻¹. -1 ; The hydrogenation catalyst comprises a support and a hydrogenation active component supported on the support; the support is alumina and / or amorphous silica-alumina, and the hydrogenation active component includes a Group VIB metal and / or a Group VIII metal.

10. The method according to claim 8, wherein, The extraction conditions include: a temperature of 80-120 ℃; a volume ratio of extraction solvent to refined gasoline of 2-6; and the extraction solvent is selected from one or more of sulfolane, N-methylpyrrolidone, diethylene glycol ether, triethylene glycol ether, tetraethylene glycol, and dimethyl sulfoxide.

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