Co-hydrotreating system and co-hydrotreating method for direct coal liquefaction oil and bio-based oil

Abstract

The invention provides a co-hydrotreating system and a co-hydrotreating method for direct coal liquefaction oil and bio-based oil. The system comprises a hydrogenation unit, a first separation unit, a co-hydrogenation treatment device and a second separation unit, the hydrogenation unit comprises a hydrofining device, a hydroisomerization device and a hydrogenation saturation device, and is used for sequentially carrying out hydrofining reaction, hydroisomerization reaction and hydrogenation saturation reaction on the direct coal liquefaction oil and hydrogen under the catalysis of a catalyst to obtain a product system; the first separation unit is used for separating a product system to obtain first naphtha, first aviation kerosene, diesel oil and transformer oil; the co-hydrogenation treatment device is used for carrying out co-hydrogenation reaction on the first naphtha and bio-based oil under the catalysis of a fourth catalyst to obtain an oil-water mixture; and the second separation unit is used for separating the oil-water mixture to obtain second naphtha, gasoline, light aromatic hydrocarbon and second aviation kerosene. By applying the system, the synergistic conversion of the direct coal liquefaction oil product and the bio-based oil can be realized.

Description

TECHNICAL FIELD

[0001] The present application relates to the technical field of coal chemical industry and comprehensive utilization of biomass energy, in particular to a coal direct liquefaction oil and bio-based oil co-hydroprocessing system and co-hydroprocessing method. BACKGROUND

[0002] Coal direct liquefaction technology is an important coal conversion approach, which converts coal into liquid fuels and chemicals through continuous hydrogenation process. Its main products include naphtha, diesel, aviation kerosene and transformer oil. However, the comprehensive utilization of coal direct liquefaction oil has not been fully developed, especially the comprehensive utilization of coal-based naphtha. The main components of coal-based naphtha are isomeric alkanes and naphthenes, which are widely used to produce motor gasoline blending oil. Although the proportion of coal-based naphtha in gasoline blending oil can be increased to some extent through catalytic reforming process, the single application still limits the potential economic value of coal-based naphtha.

[0003] On the other hand, bio-based oils (such as waste edible oil pyrolysis oil, microalgae pyrolysis oil and lignin liquefaction oil) are considered as an important part of sustainable energy due to their wide sources. However, these bio-based oils generally contain high oxygen content (8-22wt%), which makes it difficult to be directly converted into clean oil and chemicals, severely limiting its application.

[0004] The existing document (Publication No. CN104004542A) discloses a preparation method of coal-based high aromatic potential content raw material oil, which includes the following steps: (1) raw material pretreatment; (2) distillation separation; (3) slurry bed hydrocracking; (4) hydro-upgrading. This document not only can use coal tar or shale oil as raw material, but also can use coal direct liquefaction generated oil or coal oil co-refining generated oil as raw material, has wide range of applicable raw materials, can be used for producing coal-based high aromatic potential content catalytic reforming raw material, provides another way for high value-added utilization of coal-based oil, and broadens the source of high aromatic potential content catalytic reforming raw material.

[0005] The existing document (Publication No. CN108865253A) discloses a method for producing aromatic-rich oil by coal direct hydrogenation liquefaction, which includes: oil coal slurry preparation; deep hydrogenation liquefaction; atmospheric and vacuum distillation to obtain naphtha; catalytic reforming of naphtha to obtain light aromatic product. This method couples coal direct liquefaction with catalytic reforming and hydrocracking technology, and through secondary hydrocracking of diesel fraction, the purpose of maximizing production of high value-added light aromatics by coal direct liquefaction is achieved. The coal conversion rate is high, the yield of high aromatic potential content naphtha is high, the aromatic potential value is high; the yield of mixed aromatic product is high, and the content of light aromatics is high.

[0006] The prior art (Publication No. CN104927897A) discloses a coal direct liquefaction oil hydrogenation stabilization method, which comprises feeding hydrogen into coal direct liquefaction oil through pores with a mean pore size of nanometer size to obtain hydrogen-containing liquefaction oil; and feeding the hydrogen-containing liquefaction oil into a tubular fixed bed reactor in an upward flow manner to contact with a hydrogenation catalyst under liquid phase hydroprocessing conditions. The method can obtain good hydroprocessing effect, effectively reduce the sulfur and nitrogen content and the residual carbon content in the coal direct liquefaction oil, hydrogenate the polycyclic aromatic hydrocarbon part of the coal direct liquefaction oil, and reduce the gum and asphaltene content of the coal direct liquefaction oil.

[0007] The prior art (Publication No. CN119264944A) discloses an aviation kerosene and a preparation method and system thereof. The preparation method comprises the following steps: (1) distilling coal direct liquefaction oil to obtain a first light oil fraction, a first aviation kerosene fraction and a heavy oil fraction; (2) treating the heavy oil fraction by hydrocracking to obtain a second light oil fraction and a second aviation kerosene fraction; (3) mixing the first light oil fraction and the second light oil fraction, and performing alkylation reaction and hydrogenation reaction to obtain a third aviation kerosene fraction; and (4) performing hydrofining treatment on the first aviation kerosene fraction, the second aviation kerosene fraction and the third aviation kerosene fraction to obtain the aviation kerosene. The above method realizes maximized production of aviation kerosene from coal direct liquefaction oil, improves the yield of aviation kerosene, realizes high-value utilization of coal direct liquefaction oil and methanol, and reduces production cost.

[0008] At present, the treatment process of coal direct liquefaction oil mainly focuses on removing impurities (such as sulfur, nitrogen and oxygen) in the raw material, and the application of hydrogen-rich naphtha generated in the hydrogenation process is only limited to gasoline blending oil, which has a single use. Meanwhile, the treatment technology of bio-based oil is independent of the coal direct liquefaction technology system, and lacks synergistic effect, resulting in high hydrogen consumption of each treatment process, and difficulty in improving the utilization rate of bio-based oil and the product quality and economy of coal direct liquefaction oil.

[0009] Therefore, it is of great significance to research and develop a coal direct liquefaction oil and bio-based oil co-hydroprocessing system and method to reduce hydrogen consumption, improve the product quality and yield of clean oil and chemicals, and realize the synergistic conversion of coal direct liquefaction oil and bio-based oil. SUMMARY

[0010] The main purpose of the present application is to provide a coal direct liquefaction oil and bio-based oil co-hydroprocessing system and method to solve the problems of high hydrogen consumption, low product quality and yield, and inability to realize synergistic conversion with bio-based oil in the prior art coal direct liquefaction oil treatment process.

[0011] In order to achieve the above object, the present application provides a coal direct liquefaction oil and bio-based oil co-hydroprocessing system, which comprises a hydrogenation unit, a first separation unit, a co-hydroprocessing device and a second separation unit. The hydrogenation unit comprises a hydrofining device, a hydroisomerization device and a hydrogenation saturation device; the hydrofining device is used for hydrofining reaction of the coal direct liquefaction oil and hydrogen under the catalysis of a first catalyst to obtain a hydrofining product; the hydrofining device is provided with a coal direct liquefaction oil inlet, a first hydrogen inlet and a hydrofining product outlet; the hydrofining device is internally provided with a first catalyst bed layer, and the first catalyst is filled in the first catalyst bed layer; the hydroisomerization device is used for hydroisomerization reaction of the hydrofining product and hydrogen under the catalysis of a second catalyst to obtain a hydroisomerization product; the hydroisomerization device is provided with a hydrofining product inlet, a second hydrogen inlet and a hydroisomerization product outlet, and the hydrofining product inlet is in communication with the hydrofining product outlet; the hydroisomerization device is internally provided with a second catalyst bed layer, and the second catalyst is filled in the second catalyst bed layer; the hydrogenation saturation device is used for hydrogenation saturation reaction of the hydroisomerization product and hydrogen under the catalysis of a third catalyst to obtain a product system; the hydrogenation saturation device is provided with a hydroisomerization product inlet, a third hydrogen inlet and a product system outlet, and the hydroisomerization product inlet is in communication with the hydroisomerization product outlet; the hydrogenation saturation device is internally provided with a third catalyst bed layer, and the third catalyst is filled in the third catalyst bed layer; the first separation unit is used for separating the product system to obtain a first naphtha, a first aviation kerosene with a distillation range of 160-270℃, diesel oil and transformer oil with a distillation range of 290-360℃; the first separation unit is provided with a product system inlet, a first naphtha outlet, a first aviation kerosene outlet, a diesel oil outlet and a transformer oil outlet; the product system inlet is in communication with the product system outlet; the co-hydroprocessing device is used for co-hydrogenation reaction of the first naphtha and bio-based oil under the catalysis of a fourth catalyst to obtain an oil-water mixture; the co-hydroprocessing device is provided with a naphtha inlet, a bio-based oil inlet and an oil-water mixture outlet; the naphtha inlet is in communication with the first naphtha outlet; the co-hydroprocessing device is internally provided with a fourth catalyst bed layer, and the fourth catalyst is filled in the fourth catalyst bed layer; the second separation unit is used for separating the oil-water mixture to obtain a second naphtha, gasoline, light aromatic hydrocarbons and a second aviation kerosene with a distillation range of 170-240℃, and the light aromatic hydrocarbons include benzene, toluene and xylene; the second separation unit is provided with an oil-water mixture inlet, a second naphtha outlet, a gasoline outlet, a light aromatic hydrocarbon outlet and a second aviation kerosene outlet; the oil-water mixture inlet is in communication with the oil-water mixture outlet, and the second naphtha outlet is in communication with the naphtha inlet.

[0012] Further, the first separation unit comprises a first gas-liquid separation device, a second gas-liquid separation device and a first rectification device connected in sequence in a liquid phase, wherein the first gas-liquid separation device is used for separating the product system to obtain a first gas component and a hydrogenated oil product; the first gas-liquid separation device is provided with a first gas component outlet; the second gas-liquid separation device is used for separating the hydrogenated oil product to obtain a second gas component and a first oil product; the second gas-liquid separation device is provided with a second gas component outlet; the first rectification device is selected from a first rectification column, and is used for first rectifying the first oil product to obtain a first naphtha, a first aviation kerosene, diesel oil and transformer oil; the first rectification column is provided with a first naphtha outlet, a first aviation kerosene outlet, a diesel oil outlet and a transformer oil outlet from top to bottom in sequence.

[0013] Further, the first separation unit further comprises a gas separation device in communication with the first gas component outlet and used for separating the first gas component to obtain hydrogen; more preferably, the gas separation device is provided with a gas inlet and a hydrogen outlet, the gas inlet is in communication with the first gas component outlet, and the hydrogen outlet is in communication with at least one of the first hydrogen inlet, the second hydrogen inlet and the third hydrogen inlet.

[0014] Further, the second separation unit comprises an oil-water separation device and a second rectification device connected in sequence in a liquid phase, wherein the oil-water separation device is provided with a water outlet and is used for separating an oil-water mixture to obtain water and a second oil product; the second rectification device is selected from a second rectification column and is used for second rectifying the second oil product to obtain a second naphtha, gasoline, light aromatic hydrocarbon and a second aviation kerosene; the second rectification column is provided with a second naphtha outlet, a gasoline outlet, a light aromatic hydrocarbon outlet and a second aviation kerosene outlet from top to bottom in sequence.

[0015] Further, the loading amount of the first catalyst is 50-60 vol% based on the volume percentage of the hydrofining device; preferably, the first catalyst is selected from one or more of cobalt-manganese / alumina, nickel-tungsten / alumina and nickel-molybdenum / alumina.

[0016] Further, the loading amount of the second catalyst is 60-70 vol% based on the volume percentage of the hydroisomerization device; preferably, the second catalyst is selected from one or more of platinum / ZSM-5 molecular sieve, platinum / SAPO-11 molecular sieve and tungsten-molybdenum / β molecular sieve.

[0017] Further, the loading amount of the third catalyst is 55-70 vol% based on the volume percentage of the hydrogen saturation device; preferably, the third catalyst is selected from one or more of palladium / alumina, platinum / alumina and platinum-rhenium / alumina.

[0018] Further, the fourth catalyst is loaded in an amount of 45-65 vol% based on the volume percentage of the co-hydroprocessing device; preferably, the fourth catalyst is selected from one or more of the group consisting of nickel-molybdenum-tungsten / alumina, platinum-rhenium / alumina and nickel-copper / alumina.

[0019] To achieve the above-mentioned object, the present application further provides a coal direct liquefaction oil and bio-based oil co-hydroprocessing method, which is carried out by using the coal direct liquefaction oil and bio-based oil co-hydroprocessing system provided by the present application. The co-hydroprocessing method comprises the following steps: S1, carrying out a hydrofining reaction on the coal direct liquefaction oil and hydrogen under the catalysis of a first catalyst by using a hydrofining device in a hydrogenation unit to obtain a hydrofining product; carrying out a hydroisomerization reaction on the hydrofining product and hydrogen under the catalysis of a second catalyst by using a hydroisomerization device in the hydrogenation unit to obtain a hydroisomerization product; carrying out a hydro-saturation reaction on the hydroisomerization product and hydrogen under the catalysis of a third catalyst by using a hydro-saturation device in the hydrogenation unit to obtain a product system; S2, separating the product system by using a first separation unit to obtain first naphtha, first aviation kerosene with a distillation range of 160-270℃, diesel oil and transformer oil with a distillation range of 290-360℃; S3, carrying out a co-hydroprocessing reaction on the first naphtha and bio-based oil in a co-hydroprocessing device under the catalysis of a fourth catalyst to obtain an oil-water mixture; S4, separating the oil-water mixture by using a second separation unit to obtain second naphtha, gasoline, light aromatic hydrocarbons and second aviation kerosene with a distillation range of 170-240℃, the light aromatic hydrocarbons including benzene, toluene and xylene, and the second naphtha is returned to step S3 for recycling.

[0020] Further, in step S1, the hydrofining reaction is carried out at a temperature of 280-360℃, a pressure of 4-8 MPa and a volume space velocity of 0.5-2.5 h -1 .

[0021] Further, the volume ratio of hydrogen to the coal direct liquefaction oil is (1000-1500):1.

[0022] Further, the hydroisomerization reaction is carried out at a temperature of 300-400℃, a pressure of 6-10 MPa and a volume space velocity of 0.5-1.5 h -1 .

[0023] Further, the volume ratio of hydrogen to the hydrofining product is (500-700):1.

[0024] Further, the hydro-saturation reaction is carried out at a temperature of 340-420℃, a pressure of 6-10 MPa and a volume space velocity of 0.5-3 h -1 .

[0025] Further, the volume ratio of hydrogen to the hydroisomerization product is (800-1000):1.

[0026] Further, step S2 comprises: performing first gas-liquid separation on the product system by using a first gas-liquid separation device to obtain a first gas component and a hydrogenated oil product; performing second gas-liquid separation on the hydrogenated oil product by using a second gas-liquid separation device to obtain a second gas component and a first oil product; and performing first rectification on the first oil product by using a first rectification device to obtain first naphtha, first aviation kerosene, diesel oil and transformer oil.

[0027] Further, the first gas-liquid separation is performed at a temperature of 30-60°C, a pressure of 6-10 MPa and for a time of 0.1-0.5 h.

[0028] Further, the second gas-liquid separation is performed at a temperature of 40-70°C, a pressure of 0.5-3 MPa and for a time of 0.1-0.5 h.

[0029] Further, the first rectification is performed by using a first rectification column; preferably, the first rectification column has a top temperature of 100-120°C, a top pressure of 0.1-0.3 MPa, a bottom temperature of 280-340°C and a reflux ratio of (2-6):1.

[0030] Further, step S2 further comprises separating the first gas component by using a gas separation device to obtain hydrogen; preferably, the hydrogen is recycled to at least one of the hydrofining reaction, the hydroisomerization reaction and the hydro-saturation reaction in step S1.

[0031] Further, in step S3, the co-hydrogenation reaction is performed at a temperature of 300-420°C, a pressure of 0.5-1.5 MPa and a volume space velocity of 1-2.5 h -1 .

[0032] Further, the volume ratio of the first naphtha to the bio-based oil is (1.5-2):1.

[0033] Further, the bio-based oil is selected from one or more of the group consisting of waste edible oil pyrolysis oil, microalgae pyrolysis oil and lignin liquefaction oil.

[0034] Further, step S4 comprises: separating the oil-water mixture by using an oil-water separation device to obtain water and a second oil product; and performing second rectification on the second oil product by using a second rectification device to obtain second naphtha, gasoline, light aromatic hydrocarbon and second aviation kerosene.

[0035] Further, the second rectification is performed by using a second rectification column; preferably, the second rectification column has a top temperature of 30-80°C, a top pressure of 0.1-0.5 MPa, a bottom temperature of 250-320°C and a reflux ratio of (3-8):1.

[0036] By applying the technical scheme of the present application, the present application provides a coal direct liquefaction oil and bio-based oil co-hydroprocessing system, which comprises a hydrogenation unit, a first separation unit, a co-hydroprocessing device and a second separation unit. The hydrogenation unit comprises a hydrofining device, a hydroisomerization device and a hydrogenation saturation device connected in sequence. The hydrofining device is configured to effectively remove heteroatoms such as sulfur, nitrogen and oxygen in the coal direct liquefaction oil to obtain a hydrofining product. The hydroisomerization device is configured to isomerize the hydrofining product to obtain a hydroisomerization product. The hydrogenation saturation device is configured to further hydrogenate the hydroisomerization product to reduce the content of aromatic hydrocarbons and unsaturated alkanes, thereby obtaining a product system with high content of saturated hydrocarbons and naphthenes, which is convenient for subsequent separation and purification. The first separation unit is configured to separate the product system to obtain first naphtha, first aviation kerosene, diesel and transformer oil. The first separation unit is configured to separate the products in the product system and improve the quality of the products, thereby meeting the needs of different industrial applications. The co-hydroprocessing device is configured to catalyze the co-hydrogenation of naphtha and bio-based oil under the catalysis of a fourth catalyst. The excess hydrogen atoms in the naphtha can react with the oxygen atoms in the bio-based oil to achieve deep deoxygenation of the bio-based oil while generating water, thereby improving the utilization rate of the bio-based oil, reducing hydrogen consumption, and realizing the synergistic conversion of naphtha and bio-based oil. The second separation unit is configured to separate the oil-water mixture to obtain second naphtha, gasoline, light aromatics (BTX) and second aviation kerosene, thereby realizing efficient separation of the products. Meanwhile, the second naphtha can be returned to the co-hydroprocessing device as raw material through the second naphtha outlet to realize recycling, thereby improving the utilization rate of the raw material.

[0037] In summary, the co-hydroprocessing system provided by the present application can reduce hydrogen consumption, improve the yield and quality of each product, realize the synergistic conversion of coal direct liquefaction oil products and bio-based oil, thereby reducing energy waste, reducing production costs, improving the comprehensive utilization value and economic value of resources, and promoting the clean and efficient development of the energy industry. BRIEF DESCRIPTION OF DRAWINGS

[0038] The drawings accompanying the specification of the present application form a part thereof, serve to provide further understanding of the present application, and together with the description of the exemplary embodiments of the present application and the explanation thereof serve to explain the present application, and do not constitute improper limitations on the present application. In the drawings:

[0039] Figure 1 A structure schematic diagram of a coal direct liquefaction oil and bio-based oil co-hydroprocessing system in an embodiment of the present application is shown;

[0040] Figure 2 A structure schematic diagram of a co-hydroprocessing system of coal direct liquefaction oil and bio-based oil in another embodiment of the present application is shown.

[0041] In the above drawings, the following reference signs are included:

[0042] 100, hydrogenation unit; 110, hydrofining device; 111, coal direct liquefaction oil inlet; 112, first hydrogen inlet; 120, hydroisomerization device; 121, second hydrogen inlet; 130, hydrogenation saturation device; 131, third hydrogen inlet;

[0043] 200, first separation unit; 201, first aviation kerosene outlet; 202, diesel oil outlet; 203, transformer oil outlet; 210, first gas-liquid separation device; 220, second gas-liquid separation device; 230, first rectification device; 211, first gas component outlet; 221, second gas component outlet; 240, gas separation device;

[0044] 300, co-hydroprocessing device; 301, bio-based oil inlet;

[0045] 400, second separation unit; 401, gasoline outlet; 402, light aromatic hydrocarbon outlet; 403, second aviation kerosene outlet; 410, oil-water separation device; 420, second rectification device; 411, water outlet. DETAILED DESCRIPTION

[0046] It should be noted that the embodiments in the present application and the features in the embodiments can be combined with each other without conflict. The present application will be described in detail below with reference to the embodiments.

[0047] As described in the background, the existing coal direct liquefaction oil processing process has the problems of large hydrogen consumption, low product quality and yield, and inability to realize the synergistic conversion with bio-based oil. In order to solve the above technical problems, the first aspect of the present application provides a co-hydroprocessing system of coal direct liquefaction oil and bio-based oil, as shown in Figure 1As shown, the co-hydrotreating system comprises a hydrogenation unit 100, a first separation unit 200, a co-hydrotreating device 300 and a second separation unit 400. The hydrogenation unit 100 comprises a hydrofining device 110, a hydroisomerization device 120 and a hydrogenation saturation device 130; wherein the hydrofining device 110 is used for carrying out a hydrofining reaction of the coal direct liquefaction oil and hydrogen under the catalysis of a first catalyst to obtain a hydrofining product; the hydrofining device 110 is provided with a coal direct liquefaction oil inlet 111, a first hydrogen inlet 112 and a hydrofining product outlet; the hydrofining device 110 is internally provided with a first catalyst bed layer, and the first catalyst is packed in the first catalyst bed layer; the hydroisomerization device 120 is used for carrying out a hydroisomerization reaction of the hydrofining product and hydrogen under the catalysis of a second catalyst to obtain a hydroisomerization product; the hydroisomerization device 120 is provided with a hydrofining product inlet, a second hydrogen inlet 121 and a hydroisomerization product outlet, and the hydrofining product inlet is in communication with the hydrofining product outlet; the hydroisomerization device 120 is internally provided with a second catalyst bed layer, and the second catalyst is packed in the second catalyst bed layer; the hydrogenation saturation device 130 is used for carrying out a hydrogenation saturation reaction of the hydroisomerization product and hydrogen under the catalysis of a third catalyst to obtain a product system; the hydrogenation saturation device 130 is provided with a hydroisomerization product inlet, a third hydrogen inlet 131 and a product system outlet, and the hydroisomerization product inlet is in communication with the hydroisomerization product outlet; the hydrogenation saturation device 130 is internally provided with a third catalyst bed layer, and the third catalyst is packed in the third catalyst bed layer; the first separation unit 200 is used for separating the product system to obtain first naphtha, first aviation kerosene with a distillation range of 160-270°C, diesel and transformer oil with a distillation range of 290-360°C; the first separation unit 200 is provided with a product system inlet, a first naphtha outlet, a first aviation kerosene outlet 201, a diesel outlet 202 and a transformer oil outlet 203; the product system inlet is in communication with the product system outlet; the co-hydrotreating device 300 is used for carrying out a co-hydrogenation reaction of the first naphtha and bio-based oil under the catalysis of a fourth catalyst to obtain an oil-water mixture; the co-hydrotreating device 300 is provided with a naphtha inlet, a bio-based oil inlet 301 and an oil-water mixture outlet; the naphtha inlet is in communication with the first naphtha outlet; the co-hydrotreating device 300 is internally provided with a fourth catalyst bed layer, and the fourth catalyst is packed in the fourth catalyst bed layer; the second separation unit 400 is used for separating the oil-water mixture to obtain second naphtha, gasoline, light aromatic hydrocarbons and second aviation kerosene with a distillation range of 170-240°C, and the light aromatic hydrocarbons include benzene, toluene and xylene; the second separation unit 400 is provided with an oil-water mixture inlet, a second naphtha outlet, a gasoline outlet 401, a light aromatic hydrocarbon outlet 402 and a second aviation kerosene outlet 403; the oil-water mixture inlet is in communication with the oil-water mixture outlet, and the second naphtha outlet is in communication with the naphtha inlet.

[0048] The application provides a coal direct liquefaction oil and bio-based oil co-hydroprocessing system, which comprises a hydrogenation unit 100, a first separation unit 200, a co-hydroprocessing device 300 and a second separation unit 400. The hydrogenation unit 100 comprises a hydrogenation refining device 110, a hydrogenation isomerization device 120 and a hydrogenation saturation device 130 connected in sequence. The hydrogenation refining device 110 is arranged to effectively remove heteroatoms such as sulfur, nitrogen and oxygen in the coal direct liquefaction oil to obtain a hydrogenation refining product. The hydrogenation isomerization device 120 is arranged to isomerize the hydrogenation refining product to obtain a hydrogenation isomerization product. The hydrogenation saturation device 130 is arranged to further hydrogenate and saturate the hydrogenation isomerization product to reduce the content of aromatic hydrocarbons and unsaturated alkanes, so as to obtain a product system with high content of saturated hydrocarbons and naphthenes, which is convenient for subsequent separation and purification. The first separation unit 200 is arranged to separate the product system to obtain first naphtha, first aviation kerosene, diesel and transformer oil. The first separation unit 200 is arranged to separate the products in the product system and improve the product quality of the products, so as to meet the needs of different industrial applications. The co-hydroprocessing device 300 is arranged to catalyze the co-hydrogenation reaction of naphtha and bio-based oil under the catalysis of a fourth catalyst. The excess hydrogen atoms in the naphtha can react with the oxygen atoms in the bio-based oil to realize deep deoxygenation of the bio-based oil while generating water, so as to improve the utilization rate of the bio-based oil, reduce hydrogen consumption and realize the synergistic conversion of naphtha and bio-based oil. The second separation unit 400 is arranged to separate the oil-water mixture to obtain second naphtha, gasoline, light aromatic hydrocarbons (BTX) and second aviation kerosene, so as to realize the efficient separation of the products. Meanwhile, the second naphtha can be returned to the co-hydroprocessing device 300 as raw material to realize recycling, so as to improve the utilization rate of the raw material.

[0049] In summary, the above co-hydroprocessing system provided by the application can reduce hydrogen consumption, improve the yield and quality of products, realize the synergistic conversion of coal direct liquefaction oil products and bio-based oil, so as to reduce energy waste, reduce production cost, improve the comprehensive utilization value and economic value of resources, and has a positive promoting effect on promoting the clean and efficient development of the energy industry.

[0050] In a preferred embodiment, the first catalyst is loaded in an amount of 50-60 vol% based on the volume percentage of the hydrofining device 110. The loading amount of the first catalyst includes but is not limited to the above range, and the limitation of the loading amount of the first catalyst in the above range is advantageous to improve the selectivity and conversion rate of the hydrofining reaction, improve the removal efficiency of heteroatoms such as sulfur, nitrogen and oxygen, and thus is advantageous to improve the yield and purity of the hydrofining product, and at the same time is advantageous to improve the utilization rate of the first catalyst and reduce the cost. Preferably, the loading amount of the first catalyst is 55-60 vol%.

[0051] In a preferred embodiment, the first catalyst includes but is not limited to one or more of the group consisting of cobalt-manganese / alumina (Co-Mn / Al2O3), nickel-tungsten / alumina (Ni-W / Al2O3) and nickel-molybdenum / alumina (Ni-Mo / Al2O3). Compared with other types, the first catalyst of the above type has higher catalytic activity and stability, which is advantageous to improve the reaction efficiency and selectivity of the hydrofining reaction, improve the removal efficiency of heteroatoms such as sulfur, nitrogen and oxygen, and inhibit the generation of by-products, and thus is advantageous to improve the yield and purity of the hydrofining product.

[0052] In a preferred embodiment, the second catalyst is loaded in an amount of 60-70 vol% based on the volume percentage of the hydroisomerization device 120. The loading amount of the second catalyst includes but is not limited to the above range, and the limitation of the loading amount of the second catalyst in the above range is advantageous to improve the selectivity and conversion rate of the hydroisomerization reaction, promote the conversion of linear alkanes to branched alkanes, improve the structure of naphthenes, and inhibit the generation of by-products, and thus is advantageous to improve the yield and purity of the hydroisomerization product, and at the same time is advantageous to improve the utilization rate of the second catalyst and reduce the cost. Preferably, the loading amount of the second catalyst is 65-70 vol%.

[0053] In a preferred embodiment, the second catalyst includes but is not limited to one or more of the group consisting of platinum / ZSM-5 molecular sieve (Pt / ZSM-5), platinum / SAPO-11 molecular sieve (Pt / SAPO-11) and tungsten-molybdenum / β molecular sieve (W-Mo / β molecular sieve). Compared with other types, the second catalyst of the above type has higher catalytic activity and stability, which is advantageous to improve the reaction efficiency and selectivity of the hydroisomerization reaction, promote the conversion of linear alkanes to branched alkanes, improve the structure of naphthenes, and inhibit the generation of by-products, and thus is advantageous to improve the yield and purity of the hydroisomerization product.

[0054] It should be noted that the β molecular sieve in the tungsten-molybdenum / β molecular sieve catalyst is a silicoaluminate zeolite catalyst with a three-dimensional large pore structure (pore size of 0.56-0.67 nm).

[0055] In a preferred embodiment, the third catalyst is loaded in an amount of 55-70 vol% based on the volume percentage of the hydrogenation saturation device 130. The loading amount of the third catalyst, including but not limited to the above range, is limited in the above range to improve the selectivity and conversion rate of the hydrogenation saturation reaction, reduce the content of aromatic hydrocarbons and unsaturated alkanes, and inhibit the generation of by-products, thereby improving the yield and purity of the product system, while also improving the utilization of the third catalyst and reducing costs. Preferably, the third catalyst is loaded in an amount of 60-70 vol%.

[0056] In a preferred embodiment, the third catalyst includes, but is not limited to, one or more of the group consisting of palladium / alumina (Pd / Al2O3), platinum / alumina (Pt / Al2O3), and platinum-rhenium / alumina (Pt-Re / Al2O3). The third catalyst of the above-mentioned type has higher catalytic activity and stability compared to other types, which is conducive to improving the reaction efficiency and selectivity of the hydrogenation saturation reaction, reducing the content of aromatic hydrocarbons and unsaturated alkanes, and inhibiting the generation of by-products, thereby obtaining a product system with higher content of saturated hydrocarbons and naphthenes, while also improving the yield and purity of the product system.

[0057] In a preferred embodiment, the fourth catalyst is loaded in an amount of 45-65 vol% based on the volume percentage of the co-hydroprocessing device 300. The loading amount of the fourth catalyst, including but not limited to the above range, is limited in the above range to improve the selectivity and conversion rate of the co-hydrogenation reaction, promote the reaction between the excess hydrogen atoms in naphtha and the oxygen atoms in bio-oil, and inhibit the generation of by-products, thereby improving the yield and purity of the oil-water mixture, while also improving the utilization of the fourth catalyst and reducing costs. Preferably, the fourth catalyst is loaded in an amount of 45-60 vol%.

[0058] In a preferred embodiment, the fourth catalyst includes, but is not limited to, one or more of the group consisting of nickel-molybdenum-tungsten / alumina (Ni-Mo-W / Al2O3), platinum-rhenium / alumina (Pt-Re / Al2O3), and nickel-copper / alumina (Ni-Cu / Al2O3). The fourth catalyst of the above-mentioned type has higher catalytic activity and stability compared to other types, which is conducive to improving the reaction efficiency and selectivity of the co-hydrogenation reaction, promoting the reaction between the excess hydrogen atoms in naphtha and the oxygen atoms in bio-oil, and inhibiting the generation of by-products, thereby improving the yield and purity of the oil-water mixture.

[0059] In a preferred embodiment, as Figure 2As shown, the first separation unit 200 comprises a first gas-liquid separation device 210, a second gas-liquid separation device 220 and a first rectification device 230 connected in sequence, wherein the first gas-liquid separation device 210 is used to separate the product system to obtain a first gas component and a hydrogenated oil product; the first gas-liquid separation device 210 is provided with a first gas component outlet 211; the second gas-liquid separation device 220 is used to separate the hydrogenated oil product to obtain a second gas component and a first oil product; the second gas-liquid separation device 220 is provided with a second gas component outlet 221; the first rectification device 230 includes but is not limited to a first rectification column, which is used to perform first rectification on the first oil product to obtain first naphtha, first aviation kerosene, diesel and transformer oil; the first rectification column is provided with a first naphtha outlet, a first aviation kerosene outlet 201, a diesel outlet 202 and a transformer oil outlet 203 from top to bottom in sequence.

[0060] The first gas-liquid separation device 210 is provided to separate the product system obtained from the hydrogenation unit 100 to obtain a first gas component and a hydrogenated oil product, wherein the first gas component includes one or more of hydrogen, methane, carbon dioxide, carbon monoxide, hydrogen sulfide and ammonia, and the provision of the first gas-liquid separation device 210 is conducive to improving the separation efficiency of the first gas component and the hydrogenated oil product in the product system, thereby facilitating the separation of each product in the subsequent hydrogenated oil product. The second gas-liquid separation device 220 is used to separate the hydrogenated oil product to obtain a second gas component and a first oil product, the second gas component being a volatile hydrocarbon compound with a molecular weight of ≤58.12, including one or more of ethane, ethylene, propane and butane, and the provision of the second gas-liquid separation device 220 is conducive to improving the separation efficiency of the first oil product and the second gas component in the hydrogenated oil product, thereby facilitating the subsequent first rectification of the first oil product to obtain first naphtha, first aviation kerosene, diesel and transformer oil. According to the physical properties (such as boiling point) of the first naphtha, the first aviation kerosene, the diesel and the transformer oil, product outlets of different heights are provided on the side of the first rectification column, which facilitates the collection of each product obtained by the first rectification and improves the purity of the products.

[0061] In a preferred embodiment, as Figure 2As shown, the first separation unit 200 further comprises a gas separation device 240, which is in communication with the first gas component outlet 211 and is used for separating the first gas component to obtain hydrogen; preferably, the gas separation device 240 is provided with a gas inlet and a hydrogen outlet, the gas inlet is in communication with the first gas component outlet 211, and the hydrogen outlet is in communication with at least one of the first hydrogen inlet 112, the second hydrogen inlet 121 and the third hydrogen inlet 131, respectively. The provision of the gas separation device 240 is conducive to the separation of hydrogen, and at the same time, the separated hydrogen is returned to at least one of the hydrofining device 110, the hydroisomerization device 120 and the hydrogenation saturation device 130 through the hydrogen outlet to participate in the reaction to realize the recycling of hydrogen, thereby reducing the consumption of hydrogen, improving the utilization rate of hydrogen, reducing energy waste and reducing production costs.

[0062] In a preferred embodiment, as shown in Figure 2 As shown, the second separation unit 400 comprises an oil-water separation device 410 and a second rectification device 420 in sequence in liquid phase, wherein the oil-water separation device 410 is provided with a water outlet 411, the oil-water separation device 410 is used for separating the oil-water mixture to obtain water and a second oil product; the second rectification device 420 comprises but is not limited to a second rectification column, which is used for second rectification of the second oil product to obtain a second naphtha, gasoline, light aromatic hydrocarbon and a second aviation kerosene; the second rectification column is provided with a second naphtha outlet, a gasoline outlet 401, a light aromatic hydrocarbon outlet 402 and a second aviation kerosene outlet 403 in sequence from top to bottom at the side thereof.

[0063] The provision of the oil-water separation device 410 is conducive to improving the separation efficiency of the second oil product and water in the oil-water mixture, thereby facilitating the subsequent second rectification of the second oil product to obtain the second naphtha, gasoline, light aromatic hydrocarbon and second aviation kerosene. According to the physical properties (such as boiling point) of the second naphtha, gasoline, light aromatic hydrocarbon and second aviation kerosene, the product outlets of different heights are arranged at the side of the second rectification column, which facilitates the collection of the products obtained by the second rectification and improves the product purity.

[0064] It should be noted that the second rectification process also obtains heavy distillate oil with a distillation range ≥240℃, and the obtained heavy distillate oil can be collected for other subsequent treatment processes to produce clean oil products and chemicals. The subsequent treatment of the heavy distillate oil is not limited in the present application.

[0065] The second aspect of the present application also provides a method for co-hydroprocessing of coal direct liquefaction oil and bio-based oil. The co-hydroprocessing method is performed by using the co-hydroprocessing system for coal direct liquefaction oil and bio-based oil provided by the present application. The co-hydroprocessing method comprises: step S1, performing a hydrofining reaction of the coal direct liquefaction oil and hydrogen under the catalysis of a first catalyst by using the hydrofining device 110 in the hydrogenation unit 100 to obtain a hydrofining product; performing a hydroisomerization reaction of the hydrofining product and hydrogen under the catalysis of a second catalyst by using the hydroisomerization device 120 in the hydrogenation unit 100 to obtain a hydroisomerization product; performing a hydro-saturation reaction of the hydroisomerization product and hydrogen under the catalysis of a third catalyst by using the hydro-saturation device 130 in the hydrogenation unit 100 to obtain a product system; step S2, separating the product system by using the first separation unit 200 to obtain first naphtha, first aviation kerosene with a distillation range of 160-270°C, diesel oil, and transformer oil with a distillation range of 290-360°C; step S3, performing a co-hydroprocessing reaction of the first naphtha and the bio-based oil in the co-hydroprocessing device 300 under the catalysis of a fourth catalyst to obtain an oil-water mixture; step S4, separating the oil-water mixture by using the second separation unit 400 to obtain second naphtha, gasoline, light aromatic hydrocarbons, and second aviation kerosene with a distillation range of 170-240°C, wherein the light aromatic hydrocarbons include benzene, toluene, and xylene, and the second naphtha is returned to step S3 for recycling.

[0066] Compared with the single coal direct liquefaction oil treatment process in the prior art, the above coal direct liquefaction oil and bio-based oil co-hydroprocessing method provided by the present application can realize the synergistic conversion of coal direct liquefaction oil and bio-based oil. The above co-hydroprocessing method is carried out by using the above co-hydroprocessing system provided by the present application. In step S1, the hydrofining device 110 in the hydrogenation unit 100 is used to carry out the hydrofining reaction, which can effectively remove the heteroatoms such as sulfur, nitrogen and oxygen in the coal direct liquefaction oil, and obtain a hydrofining product, which provides a stable and pure raw material for the subsequent reaction; the hydroisomerization device 120 is used to carry out the hydroisomerization reaction, which can isomerize the hydrofining product to obtain a hydroisomerization product; the hydrogenation saturation device 130 is used to carry out the hydrogenation saturation reaction, which can further hydrogenate the hydroisomerization product to reduce the content of aromatic hydrocarbons and unsaturated alkanes, and obtain a product system with high content of saturated hydrocarbons and naphthenes, which is convenient for subsequent separation and purification. In step S2, the first separation unit 200 is used to separate the product system, which can obtain first naphtha, first aviation kerosene, diesel and transformer oil, realize the separation of each product in the product system, improve the purity and product quality of the above products, and thus meet the needs of different industrial applications. In step S3, the co-hydroprocessing device 300 is used to make the naphtha and the bio-based oil undergo co-hydrogenation reaction under the catalysis of the fourth catalyst, the excess hydrogen atoms in the naphtha can react with the oxygen atoms in the bio-based oil to generate water and realize deep deoxygenation of the bio-based oil, thereby improving the utilization rate of the bio-based oil, reducing the hydrogen consumption, and realizing the synergistic conversion of the naphtha and the bio-based oil. In step S4, the second separation unit 400 is used to separate the oil-water mixture, which can obtain second naphtha, gasoline, light aromatic hydrocarbons (BTX) and second aviation kerosene, realize the efficient separation of the above products, improve the purity and product quality, and at the same time, return the second naphtha to step S3 as a raw material to participate in the co-hydrogenation reaction to realize the recycling, thereby improving the utilization rate of the raw material.

[0067] In summary, the above coal direct liquefaction oil and bio-based oil co-hydroprocessing method provided by the present application can reduce the hydrogen consumption, improve the yield and quality of each product, realize the synergistic conversion of coal direct liquefaction oil products and bio-based oil, thereby reducing energy waste, reducing production cost, improving the comprehensive utilization value and economic value of resources, and having a positive promoting effect on the clean and efficient development of the energy industry.

[0068] In a preferred embodiment, in step S1, the temperature of the hydrofining reaction is 280-360℃, the pressure is 4-8MPa, the volume space velocity is 0.5-2.5h -1The volume space velocity of the hydrofining reaction refers to the ratio of the volume flow rate of the coal direct liquefaction oil to the volume of the first catalyst per unit time. The temperature, pressure, and volume space velocity of the hydrofining reaction include but are not limited to the above ranges, and limiting them within the above ranges is beneficial to improving the reaction efficiency of the hydrofining reaction, improving the removal efficiency of the heteroatoms such as sulfur, nitrogen, and oxygen, and inhibiting the generation of byproducts, thereby being beneficial to improving the yield and purity of the hydrofining product. Preferably, the temperature of the hydrofining reaction is 300-340°C, the pressure is 5-7 MPa, and the volume space velocity is 1-2 h -1 .

[0069] In order to further improve the selectivity and conversion rate of the hydrofining reaction, further improve the removal efficiency of the heteroatoms such as sulfur, nitrogen, and oxygen, and inhibit the generation of byproducts, thereby further improving the yield and purity of the hydrofining product, in a preferred embodiment, the volume ratio of hydrogen to the coal direct liquefaction oil is (1000-1500): 1.

[0070] In a preferred embodiment, the temperature of the hydroisomerization reaction is 300-400°C, the pressure is 6-10 MPa, and the volume space velocity is 0.5-1.5 h -1 . The volume space velocity of the hydroisomerization reaction refers to the ratio of the volume flow rate of the hydrofining product to the volume of the second catalyst per unit time. The temperature, pressure, and volume space velocity of the hydroisomerization reaction include but are not limited to the above ranges, and limiting them within the above ranges is beneficial to improving the reaction efficiency of the hydroisomerization reaction, promoting the conversion of linear alkanes to branched alkanes, improving the structure of naphthenes, and inhibiting the generation of byproducts, thereby being beneficial to improving the yield and purity of the hydroisomerization product. Preferably, the temperature of the hydroisomerization reaction is 330-380°C, the pressure is 7-9 MPa, and the volume space velocity is 0.5-1 h -1 .

[0071] In order to further improve the selectivity and conversion rate of the hydroisomerization reaction, further promote the conversion of linear alkanes to branched alkanes, improve the structure of naphthenes, and inhibit the generation of byproducts, thereby further improve the yield and purity of the hydroisomerization product, in a preferred embodiment, the volume ratio of hydrogen to the hydrofining product is (500-700): 1.

[0072] In a preferred embodiment, the temperature of the hydroisomerization reaction is 300-400°C, the pressure is 6-10 MPa, and the volume space velocity is 0.5-1.5 h -1The volume space velocity of the hydrogenation saturation reaction refers to the ratio of the volume flow rate of the hydrogenation isomerization product per unit time to the volume of the third catalyst. The temperature, pressure and volume space velocity of the hydrogenation saturation reaction include but are not limited to the above ranges, which are limited in the above ranges to improve the reaction efficiency of the hydrogenation saturation reaction, reduce the content of aromatic hydrocarbons and unsaturated alkanes, inhibit the generation of by-products, thereby facilitating to obtain a product system with higher content of saturated hydrocarbons and naphthenes, and meanwhile, improving the yield and purity of the product system. Preferably, the temperature of the hydrogenation saturation reaction is 360-400℃, the pressure is 7-9 MPa, and the volume space velocity is 1-2.5 h -1 .

[0073] In order to further improve the selectivity and conversion rate of the hydrogenation saturation reaction, further reduce the content of aromatic hydrocarbons and unsaturated alkanes, inhibit the generation of by-products, thereby further improve the yield and purity of the product system, in a preferred embodiment, the volume ratio of hydrogen to hydrogenation isomerization product is (800-1000):1.

[0074] In a preferred embodiment, step S2 comprises: using the first gas-liquid separation device 210 to perform first gas-liquid separation on the product system to obtain a first gas component and a hydrogenated oil product; using the second gas-liquid separation device 220 to perform second gas-liquid separation on the hydrogenated oil product to obtain a second gas component and a first oil product; and using the first rectification device 230 to perform first rectification on the first oil product to obtain first naphtha, first aviation kerosene, diesel and transformer oil. Compared with other methods, using the above method is advantageous to improve the separation efficiency of each material, and obtain first naphtha, first aviation kerosene, diesel and transformer oil with higher purity, thereby improving product quality and production efficiency.

[0075] In order to improve the separation efficiency of the first gas-liquid separation and obtain the first gas component and the hydrogenated oil product with higher purity, in a preferred embodiment, the temperature of the first gas-liquid separation is 30-60℃, the pressure is 6-10 MPa, and the time is 0.1-0.5 h.

[0076] In order to improve the separation efficiency of the second gas-liquid separation and obtain the second gas component and the first oil product with higher purity, in a preferred embodiment, the temperature of the second gas-liquid separation is 40-70℃, the pressure is 0.5-3 MPa, and the time is 0.1-0.5 h.

[0077] In a preferred embodiment, the first rectification is performed by using a first rectification column, preferably, the overhead temperature of the first rectification column is 100-120℃, the overhead pressure is 0.1-0.3MPa, the bottom temperature is 280-340℃, and the reflux ratio is (2-6):1. Compared with other ranges, limiting the above process parameters of the first rectification column in the above ranges is advantageous to improve the separation efficiency of the first rectification, improve the purity and yield of the first naphtha, the first aviation kerosene, diesel and transformer oil, and is also advantageous to reduce energy consumption and improve production efficiency.

[0078] In a preferred embodiment, step S2 further comprises separating the first gaseous component by using the gas separation device 240 to obtain hydrogen, and preferably, the hydrogen is recycled to at least one of the hydrofining reaction, the hydroisomerization reaction and the hydro-saturation reaction in step S1. Separating the first gaseous component by using the gas separation device 240 is advantageous to obtain hydrogen, and the separated hydrogen is returned to at least one of the hydrofining reaction, the hydroisomerization reaction and the hydro-saturation reaction in step S1 to participate in the reaction to realize the recycling of hydrogen, thereby being advantageous to reduce the consumption of hydrogen, improve the utilization rate of hydrogen, reduce energy waste and reduce production cost.

[0079] In a preferred embodiment, in step S3, the temperature of the co-hydrogenation reaction is 300-420℃, the pressure is 0.5-1.5MPa, and the volume space velocity is 1-2.5h -1 The volume space velocity of the co-hydrogenation reaction refers to the ratio of the total volume flow rate of the first naphtha and the bio-based oil to the volume of the fourth catalyst per unit time. The temperature, pressure and volume space velocity of the co-hydrogenation reaction include but are not limited to the above ranges, and limiting them in the above ranges is advantageous to improve the reaction efficiency of the co-hydrogenation reaction, promote the reaction between the excess hydrogen atoms in the naphtha and the oxygen atoms in the bio-oil, and inhibit the generation of by-products, thereby being advantageous to improve the yield and purity of the oil-water mixture and facilitate subsequent separation and purification. Preferably, the temperature of the co-hydrogenation reaction is 340-400℃, the pressure is 1-1.5MPa, and the volume space velocity is 1.5-2.5h -1 .

[0080] In order to further improve the selectivity and conversion rate of the co-hydrogenation reaction, further promote the reaction between the excess hydrogen atoms in the naphtha and the oxygen atoms in the bio-oil, and inhibit the generation of by-products, thereby further improving the yield and purity of the oil-water mixture, in a preferred embodiment, the volume ratio of the first naphtha to the bio-based oil is (1.5-2):1. Specifically, the volume ratio of the first naphtha to the bio-based oil can be 1.5:1, 1.6:1, 1.7:1, 1.8:1, 1.9:1 or 2:1.

[0081] Waste cooking oil pyrolysis oil is an oil product obtained by converting waste cooking oil through a pyrolysis process, and the oxygen content thereof is 5-20wt%. Microalgae pyrolysis oil is a bio-oil extracted from microalgae through a pyrolysis technology, and the oxygen content thereof is 10-30wt%. Lignin liquefaction oil is a product obtained by liquefying lignin under certain conditions, and the oxygen content thereof is 15-35wt%. In a preferred embodiment, the bio-based oil includes one or more of the group consisting of waste cooking oil pyrolysis oil, microalgae pyrolysis oil and lignin liquefaction oil. The bio-based oil of the above-mentioned types is easy to obtain, and has a more suitable oxygen content, which is conducive to reacting with hydrogen atoms in naphtha, thereby facilitating the reduction of the oxygen content of the bio-based oil, realizing the co-conversion with naphtha, and obtaining clean oil products and chemicals such as second naphtha, gasoline, light aromatic hydrocarbons and second aviation kerosene.

[0082] In a preferred embodiment, step S4 includes: separating the oil-water mixture by using an oil-water separation device 410 to obtain water and a second oil product; and performing second rectification on the second oil product by using a second rectification device 420 to obtain second naphtha, gasoline, light aromatic hydrocarbons and second aviation kerosene. Compared with other methods, the use of the above-mentioned method is conducive to improving the separation efficiency of each material, obtaining second naphtha, gasoline, light aromatic hydrocarbons and second aviation kerosene with higher purity, and thereby improving product quality and production efficiency.

[0083] In a preferred embodiment, the second rectification is performed by using a second rectification column, and the top temperature of the second rectification column is 30-80℃, the top pressure is 0.1-0.5MPa, the bottom temperature is 250-320℃, and the reflux ratio is (3-8):1. Compared with other ranges, limiting the above-mentioned process parameters of the second rectification column within the above-mentioned ranges is conducive to improving the separation efficiency of the second rectification, improving the purity and yield of second naphtha, gasoline, light aromatic hydrocarbons and second aviation kerosene, and simultaneously reducing energy consumption and improving production efficiency.

[0084] In order to further improve the separation efficiency of the second rectification, further improve the purity and yield of second naphtha, gasoline, light aromatic hydrocarbons and second aviation kerosene, and further reduce energy consumption and improve production efficiency, preferably, the top temperature of the second rectification column is 40-60℃, the top pressure is 0.1-0.3MPa, the bottom temperature is 260-300℃, and the reflux ratio is (4-6):1.

[0085] The application will be further described in detail below in combination with specific examples, which cannot be understood as limiting the scope of the application claimed.

[0086] It should be noted that the hydrogenation isomerization products in all examples and comparative examples of the application are tested as follows:

[0087] (1) Freezing point: according to the method recorded in national standard GB / T 2430-2008 "Determination of freezing point of aviation fuel", the freezing point was measured by using an automatic freezing point tester;

[0088] (2) Flash point: according to the method recorded in national standard GB / T 21789-2008 "Determination of flash point of petroleum products and other liquids - Abel closed cup method", the flash point was measured by using an automatic closed cup flash point tester;

[0089] (3) Density: according to the method recorded in national standard GB / T 1884-2000 "Determination of density of crude petroleum and liquid petroleum products - Laboratory method (hydrometer method)", the density was measured by using a density tester;

[0090] (4) Calorific value: according to the method recorded in national standard GB / T 384-1981 "Determination of calorific value of petroleum products", the calorific value was measured by using a calorific value tester;

[0091] (5) Viscosity: according to the method recorded in national standard GB / T 265-1988 "Determination of kinematic viscosity and calculation of dynamic viscosity of petroleum products", the viscosity was measured by using a kinematic viscosity tester.

[0092] Example 1

[0093] A method for co-hydroprocessing of coal direct liquefaction oil and bio-based oil was carried out by using a co-hydroprocessing system of coal direct liquefaction oil and bio-based oil as shown in Figure 2 , which specifically included the following steps:

[0094] (1) preparing coal direct liquefaction oil raw material, the specific composition of which was shown in Table 1; preparing bio-based oil raw material, the specific types and compositions of which were shown in Table 2;

[0095] (2) Co-Mo / Al2O3 catalyst bed was arranged inside the hydrofining reactor (Beijing Kunlun Yongtai Technology Co., Ltd., KLCP2010), and the loading amount of Co-Mo / Al2O3 catalyst was 57 vol% based on the volume percentage content of the hydrofining reactor; hydrogen and the coal direct liquefaction oil prepared in step (1) were introduced into the hydrofining reactor, and hydrofining reaction was carried out under the conditions of 300°C and 5 MPa to obtain a hydrofining product; wherein the volume space velocity of the hydrofining reaction was 2h -1 -1, and the volume ratio of hydrogen to coal direct liquefaction oil was 1300:1; the sulfur content in the prepared hydrofining product was 0.03 wt%, the nitrogen content was 0.024 wt%, and the oxygen content was 0.048 wt%;

[0096] (3) A Pt / ZSM-5 catalyst bed is set inside the hydroisomerization reactor (Beijing Kunlun Yongtai Technology Co., Ltd., catalytic cracking fluidized bed reactor). The loading amount of Pt / ZSM-5 catalyst is 66 vol% based on the volume percentage of the hydroisomerization reactor. Hydrogen gas and the hydrorefined product obtained in step (2) are introduced into the hydroisomerization reactor, and the hydroisomerization reaction is carried out at 330℃ and 7MPa to obtain the hydroisomerized product. The volume hourly space velocity of the hydroisomerization reaction is 1 h⁻¹. -1 The volume ratio of hydrogen to the hydrorefined product was 600:1; the resulting hydroisomer had a freezing point of -40℃, a flash point of 65℃, and a density of 0.81 g / cm³. 3 Its calorific value is 43.5 MJ / kg, and its viscosity is 2.5 mm. 2 / s;

[0097] (4) A Pd / Al2O3 catalyst bed is set inside the hydrogenation saturation reactor (Beijing Kunlun Yongtai Technology Co., Ltd., catalytic cracking fluidized bed reactor). The loading amount of Pd / Al2O3 catalyst is 55 vol% based on the volume percentage of the hydrogenation saturation reactor. Hydrogen gas and the hydroisomer obtained in step (3) are introduced into the hydrogenation saturation reactor, and the hydrogenation saturation reaction is carried out at 360℃ and 7MPa to obtain the product system. The volume hourly space velocity of the hydrogenation saturation reaction is 2 h⁻¹. -1 The volume ratio of hydrogen to the hydroisomer was 900:1; the product system contained 10 wt% aromatics and 70 wt% cycloalkanes by weight.

[0098] (5) Under the conditions of 40℃ and 7MPa, the product system obtained in step (4) is subjected to gas-liquid separation for 0.2h using the first gas-liquid separation device 210 (Beijing Kunlun Yongtai Technology Co., Ltd., catalytic distillation device) to obtain the first gas component and hydrogenated oil. The first gas component includes hydrogen, ammonia, hydrogen sulfide and carbon dioxide. The first gas component is separated using the gas separation device 240 (Beijing Kunlun Yongtai Technology Co., Ltd., catalytic distillation device) to obtain hydrogen. The obtained hydrogen is returned to steps (2), (3) and (4) for recycling.

[0099] (6) At 50℃ and 1MPa, the hydrogenated oil was subjected to gas-liquid separation for 0.3h using a second gas-liquid separation device 220 (Beijing Kunlun Yongtai Technology Co., Ltd., catalytic distillation device) to obtain a second gas component and a first oil product, wherein the second gas component includes ethane, propane and butane;

[0100] (7) The first oil product is subjected to rectification by using a first rectification column to obtain first naphtha, first aviation kerosene, diesel oil and transformer oil, wherein the overhead temperature of the first rectification column is 110°C, the overhead pressure is 0.2 MPa, the bottom temperature is 320°C, the reflux ratio is 3:1, the yield of the first naphtha is 21%, the yield of the first aviation kerosene is 27%, the yield of the diesel oil is 33%, and the yield of the transformer oil is 19%;

[0101] (8) A Ni-Mo-W / Al2O3 catalyst bed is arranged inside a co-hydrogenation reactor (Kunlun Yongtai Technology Co., Ltd., KLL-2201pc), and the loading amount of the Ni-Mo-W / Al2O3 catalyst is 55 vol% based on the volume percentage in the co-hydrogenation reactor; hydrogen, the first naphtha obtained in step (7) and the bio-based oil prepared in step (1) are introduced into the co-hydrogenation reactor, and co-hydrogenation reaction is carried out at 360°C and 1.2 MPa to obtain an oil-water mixture; wherein the volume space velocity of the co-hydrogenation reaction is 2h -1 , the ratio of the volume of hydrogen to the total volume of the first naphtha and the bio-based oil is 1200:1, and the volume ratio of the first naphtha to the bio-based oil is 1.8:1;

[0102] (9) The oil-water separation device 410 is used to separate the oil-water mixture obtained in step (8) to obtain water and a second oil product; the second oil product is subjected to rectification by using a second rectification column to obtain second naphtha, gasoline, BTX and second aviation kerosene, and the obtained second naphtha is returned to step (8) for recycling; wherein the overhead temperature of the second rectification column is 60°C, the overhead pressure is 0.3 MPa, the bottom temperature is 320°C, the reflux ratio is 4:1, the yield of the second naphtha is 12%, the yield of the gasoline is 18%, the yield of the BTX is 10%, and the yield of the second aviation kerosene is 50%.

[0103] Example 2

[0104] The difference from example 1 is that:

[0105] In step (1), the composition of the coal direct liquefaction oil raw material is different from that of example 1, and the specific composition is shown in Table 1; the type and composition of the bio-based oil raw material are different from those of example 1, and the specific composition is shown in Table 2;

[0106] In step (2), the temperature of the hydrofining reaction is 340°C, the pressure is 7 MPa, and the catalyst is a Ni-W / Al2O3 catalyst; the sulfur content in the hydrofining product prepared in step (2) is 0.014 wt%, the nitrogen content is 0.02 wt%, and the oxygen content is 0.05 wt%;

[0107] The temperature of the hydrogenation isomerization reaction in step (3) is 380℃, the pressure is 9 MPa, and the catalyst is a Pt / SAPO-11 catalyst; the freezing point of the hydrogenation isomerization product prepared in step (3) is -48℃, the flash point is 68℃, the density is 0.79 g / cm 3 , the calorific value is 44.2 MJ / kg, and the viscosity is 2.0 mm 2 / s;

[0108] The temperature of the hydrogenation saturation reaction in step (4) is 400℃, the pressure is 9 MPa, and the catalyst is a Pt-Re / Al2O3 catalyst; the weight percentage content of aromatic hydrocarbons in the product system prepared in step (4) is 8 wt%, and the weight percentage content of naphthenes is 80 wt%;

[0109] The temperature of the co-hydrogenation reaction in step (8) is 400℃, the pressure is 1.5 MPa, the catalyst is a Pt-Re / Al2O3 catalyst, and the volume ratio of the first naphtha to the bio-based oil prepared in step (1) is 2:1;

[0110] The remaining steps are the same as in Example 1.

[0111] The yield of the first naphtha, the first aviation kerosene, the diesel, and the transformer oil obtained by rectification in step (7) and the yield of the second naphtha, the gasoline, the BTX, and the second aviation kerosene obtained by rectification in step (9) are shown in Table 3.

[0112] Example 3

[0113] The difference from Example 1 is that:

[0114] In step (1), the composition of the coal direct liquefaction oil raw material is different from that in Example 1, and is specifically shown in Table 1; the types and compositions of the bio-based oil raw material are different from those in Example 1, and are specifically shown in Table 2.

[0115] The temperature of the hydrofining reaction in step (2) is 280℃, the pressure is 4 MPa, and the catalyst is a Ni-Mo / Al2O3 catalyst; the sulfur content in the hydrofining product prepared in step (2) is 0.01 wt%, the nitrogen content is 0.015 wt%, and the oxygen content is 0.1 wt%;

[0116] The temperature of the hydrogenation isomerization reaction in step (3) is 300℃, the pressure is 6 MPa, and the catalyst is a W-Mo / β molecular sieve catalyst; the freezing point of the hydrogenation isomerization product prepared in step (3) is -50℃, the flash point is 70℃, the density is 0.78 g / cm 3 , the calorific value is 44.8 MJ / kg, and the viscosity is 1.8 mm 2 / s;

[0117] In step (4), the temperature of the hydrogenation saturation reaction is 340℃, the pressure is 6 MPa, and the catalyst is a Pt / γ-Al2O3 catalyst; the weight percentage content of aromatic hydrocarbons in the product system prepared in step (4) is 5wt%, and the weight percentage content of naphthenes is 85wt%;

[0118] In step (8), the temperature of the co-hydrogenation reaction is 300℃, the pressure is 0.5 MPa, the catalyst is a Ni-Cu / Al2O3 catalyst, and the volume ratio of the first naphtha to the bio-based oil prepared in step (1) is 1.5:1;

[0119] The remaining steps are the same as in Example 1.

[0120] The yields of the first naphtha, the first aviation kerosene, the diesel, and the transformer oil obtained by rectification in step (7), and the yields of the second naphtha, the gasoline, the BTX, and the second aviation kerosene obtained by rectification in step (9) are shown in Table 3.

[0121] The specific compositions of the coal direct liquefaction oil raw materials in Examples 1 to 3 are shown in Table 1; the specific types and compositions of the bio-based oil raw materials in Examples 1 to 3 are shown in Table 2.

[0122] Table 1

[0123]

[0124] Table 2

[0125] Class Distillation range (°C) Oxygen content (wt%) Example 1 Waste cooking oil pyrolysis oil 100～300 18 Example 2 Microalgal pyrolysis oil 60～260 22 Example 3 Lignin liquefaction oil 120～240 8

[0126] Example 4

[0127] The difference from Example 1 is that in step (2), the temperature of the hydrofining reaction is 360℃, the pressure is 8 MPa, and the sulfur content in the hydrofining product prepared in step (2) is 0.020wt%, the nitrogen content is 0.018wt%, and the oxygen content is 0.041wt%; the remaining steps are the same as in Example 1.

[0128] Example 5

[0129] The difference from Example 1 is that in step (2), the temperature of the hydrofining reaction is 260℃, the pressure is 3 MPa, and the sulfur content in the hydrofining product prepared in step (2) is 0.035wt%, the nitrogen content is 0.027wt%, and the oxygen content is 0.051wt%; the remaining steps are the same as in Example 1.

[0130] Example 6

[0131] The difference from Example 1 is that in step (3), the temperature of the hydroisomerization reaction is 400℃, the pressure is 10 MPa, and the freezing point of the hydroisomerization product prepared in step (3) is -41℃, the flash point is 63℃, the density is 0.79 g / cm 3 , the calorific value is 44.2 MJ / kg, and the viscosity is 2.3 mm 2 / s; the remaining steps are the same as in Example 1.

[0132] Example 7

[0133] The difference from Example 1 is that in step (3), the temperature of the hydroisomerization reaction is 280℃, the pressure is 4 MPa, and the freezing point of the hydroisomerization product prepared in step (3) is -23℃, the flash point is 80℃, the density is 0.86 g / cm 3 , the calorific value is 48.6 MJ / kg, and the viscosity is 3.6 mm 2 / s; the remaining steps are the same as in Example 1.

[0134] Example 8

[0135] The difference from Example 1 is that in step (4), the temperature of the hydrogenation saturation reaction is 420℃, the pressure is 10 MPa, and the weight percentage of aromatics in the product system prepared in step (4) is 3wt%, the weight percentage of naphthenes is 77wt%; the remaining steps are the same as in Example 1.

[0136] Example 9

[0137] The difference from Example 1 is that in step (4), the temperature of the hydrogenation saturation reaction is 300℃, the pressure is 4 MPa, and the weight percentage of aromatics in the product system prepared in step (4) is 16wt%, the weight percentage of naphthenes is 64wt%; the remaining steps are the same as in Example 1.

[0138] Example 10

[0139] The difference from Example 1 is that in step (8), the volume ratio of the first naphtha to the bio-based oil is 1.5:1, and the volume space velocity of the co-hydrogenation reaction is 1 h -1 ; the remaining steps are the same as in Example 1.

[0140] Example 11

[0141] The difference from Example 1 is that in step (8), the volume ratio of the first naphtha to the bio-based oil is 2:1, and the volume space velocity of the co-hydrogenation reaction is 2.5 h -1 ; the remaining steps are the same as in Example 1.

[0142] Example 12

[0143] The difference from example 1 is that in step (8), the volume ratio of the first naphtha to the bio-based oil is 1.2:1, and the volume space velocity of the co-hydrogenation reaction is 2.6h -1 The remaining steps are the same as example 1.

[0144] Example 13

[0145] The difference from example 1 is that in step (8), the temperature of the co-hydrogenation reaction is 300℃, and the pressure is 0.5MPa; the remaining steps are the same as example 1.

[0146] Example 14

[0147] The difference from example 1 is that in step (8), the temperature of the co-hydrogenation reaction is 420℃, and the pressure is 1.5MPa; the remaining steps are the same as example 1.

[0148] Example 15

[0149] The difference from example 1 is that in step (8), the temperature of the co-hydrogenation reaction is 280℃, and the pressure is 0.3MPa; the remaining steps are the same as example 1.

[0150] Comparative example 1

[0151] The difference from example 1 is that in step (1), the bio-based oil raw material is not prepared, and steps (8) and (9) are omitted, and the remaining steps are the same as example 1.

[0152] Comparative example 1 cannot achieve the synergistic conversion of coal direct liquefaction oil and bio-based oil because bio-based oil is not introduced, so it cannot obtain the second naphtha, gasoline, BTX and second aviation kerosene.

[0153] The yield of the first naphtha, the first aviation kerosene, the diesel and the transformer oil obtained by step (7) in the above examples 1 to 15 and comparative example 1 of the present application, and the yield of the second naphtha, the gasoline, the BTX and the second aviation kerosene obtained by step (9) in examples 1 to 15 are shown in Table 3.

[0154] Table 3

[0155]

[0156] From the above description, it can be seen that the above examples of the present application achieve the following technical effects:

[0157] It can be seen from comparative example 1 and example 1 that comparative example 1 cannot realize the synergistic conversion of the coal direct liquefaction oil and the bio-based oil because the bio-based oil is not introduced, and thus the second naphtha, the gasoline, the BTX and the second aviation kerosene cannot be obtained. It can be known that the synergistic conversion of the coal direct liquefaction oil product and the bio-based oil can be realized by using the above-mentioned coal direct liquefaction oil and bio-based oil co-hydroprocessing system provided in the application, the energy waste can be reduced, the production cost can be reduced, and the comprehensive utilization value and the economic value of the resources can be improved.

[0158] It can be seen from comparative example 1, example 4 and example 5 that the yield of each product except the transformer oil in example 5 is lower than that in example 1 and example 4. The yield of the transformer oil in example 5 is higher because the product at the lower end of the first rectifying tower is the transformer oil, that is, the remaining oil product obtained by the first rectification can be used as the transformer oil, and thus the yield of the transformer oil in example 5 is higher than that in example 1 and example 4. It can be known that compared with other ranges, limiting the temperature and the pressure of the hydrofining reaction in the above-mentioned ranges is beneficial to improving the reaction efficiency of the hydrofining reaction, improving the removal efficiency of the heteroatoms such as sulfur, nitrogen and oxygen, and inhibiting the generation of by-products, so as to be beneficial to improving the yield and the purity of the hydrofining product, and further beneficial to improving the yield and the purity of the above-mentioned products.

[0159] It can be seen from comparative example 1, example 6 and example 7 that the yield of each product except the transformer oil in example 7 is lower than that in example 1 and example 6. The yield of the transformer oil in example 7 is higher because the product at the lower end of the first rectifying tower is the transformer oil, that is, the remaining oil product obtained by the first rectification can be used as the transformer oil, and thus the yield of the transformer oil in example 7 is higher than that in example 1 and example 6. It can be known that compared with other ranges, limiting the temperature and the pressure of the hydroisomerization reaction in the above-mentioned ranges is beneficial to improving the reaction efficiency of the hydroisomerization reaction, promoting the conversion of the linear alkanes to the branched alkanes, improving the structure of the naphthenes, and inhibiting the generation of by-products, so as to be beneficial to improving the yield and the purity of the hydroisomerization product, and further beneficial to improving the yield and the purity of the above-mentioned products.

[0160] As can be seen from comparative examples 1, 8 and 9, the yield of each product except transformer oil in example 9 is lower than that in examples 1 and 8, and the yield of transformer oil in example 9 is higher because the product at the lower end of the first rectifying tower is transformer oil, that is, the remaining oil product obtained by the first rectification can be used as transformer oil, so the yield of transformer oil in example 9 is higher than that in examples 1 and 8. Therefore, compared with other ranges, limiting the temperature and pressure of the hydrogenation saturation reaction in the above ranges is beneficial to improve the reaction efficiency of the hydrogenation saturation reaction, reduce the content of aromatic hydrocarbons and unsaturated alkanes, and inhibit the generation of by-products, so as to obtain a product system with higher content of saturated hydrocarbons and naphthenes, and further improve the yield and purity of each product.

[0161] As can be seen from comparative examples 1, 10 to 12, compared with other ranges, limiting the volume ratio of the first naphtha to the bio-based oil and the volume space velocity of the co-hydrogenation reaction in the above ranges is beneficial to improve the selectivity and conversion rate of the co-hydrogenation reaction, promote the reaction between the excess hydrogen atoms in naphtha and the oxygen atoms in bio-oil, and at the same time inhibit the generation of by-products, so as to improve the yield and purity of the second naphtha, gasoline, BTX and second aviation kerosene.

[0162] As can be seen from comparative examples 1, 13 to 15, compared with other ranges, limiting the temperature and pressure of the co-hydrogenation reaction in the above ranges is beneficial to improve the reaction efficiency of the co-hydrogenation reaction, promote the reaction between the excess hydrogen atoms in naphtha and the oxygen atoms in bio-oil, and at the same time inhibit the generation of by-products, so as to improve the yield and purity of the second naphtha, gasoline, BTX and second aviation kerosene.

[0163] It should be noted that the terms "first", "second" and the like in the specification and claims of the present application are used to distinguish similar objects, and do not necessarily have to describe a specific order or sequence. It should be understood that the terms thus used can be interchanged under appropriate circumstances, so that the embodiments of the present application described herein can be implemented in an order other than that described herein.

[0164] The above only describes the preferred embodiments of the present application and is not intended to limit the present application. For those skilled in the art, the present application can have various modifications and changes. Any modification, equivalent replacement, improvement, etc. made within the spirit and principles of the present application shall be included in the protection scope of the present application.

Claims

1. A system for co-hydrogenation of coal direct liquefaction oil and bio-based oil, characterized in that, The co-hydrogenation processing system includes: Hydrogenation unit (100), comprising: A hydrorefining unit (110) is used to react the direct coal liquefaction oil with hydrogen under the catalysis of a first catalyst to obtain a hydrorefined product; the hydrorefining unit (110) is provided with a direct coal liquefaction oil inlet (111), a first hydrogen inlet (112) and a hydrorefined product outlet; a first catalyst bed is provided inside the hydrorefining unit (110), and the first catalyst is packed in the first catalyst bed; A hydroisomerization device (120) is used to react the hydrorefined product with hydrogen under the catalysis of a second catalyst to obtain a hydroisomerized product; the hydroisomerization device (120) is provided with a hydrorefined product inlet, a second hydrogen inlet (121) and a hydroisomerized product outlet, the hydrorefined product inlet and the hydrorefined product outlet being connected; a second catalyst bed is provided inside the hydroisomerization device (120), and the second catalyst is packed in the second catalyst bed; A hydrogenation saturation unit (130) is used to cause the hydrogenation isomer product and hydrogen to undergo a hydrogenation saturation reaction under the catalysis of a third catalyst to obtain a product system. The hydrogenation saturation unit (130) is provided with a hydrogenation isomer product inlet, a third hydrogen inlet (131), and a product system outlet. The hydrogenation isomer product inlet is connected to the hydrogenation isomer product outlet. A third catalyst bed is provided inside the hydrogenation saturation unit (130), and the third catalyst is packed in the third catalyst bed. A first separation unit (200) is used to separate the product system to obtain first naphtha, first aviation kerosene with a distillation range of 160-270℃, diesel oil, and transformer oil with a distillation range of 290-360℃. The first separation unit (200) is provided with a product system inlet, a first naphtha outlet, a first aviation kerosene outlet (201), a diesel oil outlet (202), and a transformer oil outlet (203). The product system inlet is connected to the product system outlet. A co-hydrogenation treatment device (300) is used to co-hydrogenate the first naphtha and the bio-based oil under the catalysis of a fourth catalyst to obtain an oil-water mixture; the co-hydrogenation treatment device (300) is provided with a naphtha inlet, a bio-based oil inlet (301) and an oil-water mixture outlet; the naphtha inlet is connected to the first naphtha outlet; the co-hydrogenation treatment device (300) is provided with a fourth catalyst bed, and the fourth catalyst is packed in the fourth catalyst bed; The second separation unit (400) is used to separate the oil-water mixture to obtain second naphtha, gasoline, light aromatics and second aviation kerosene with a distillation range of 170-240°C. The light aromatics include benzene, toluene and xylene. The second separation unit (400) is provided with an oil-water mixture inlet, a second naphtha outlet, a gasoline outlet (401), a light aromatics outlet (402) and a second aviation kerosene outlet (403). The oil-water mixture inlet is connected to the oil-water mixture outlet, and the second naphtha outlet is connected to the naphtha inlet.

2. The coal direct liquefaction oil and bio-based oil co-hydrogenation treatment system according to claim 1, characterized in that, The first separation unit (200) includes a first gas-liquid separation device (210), a second gas-liquid separation device (220), and a first distillation device (230) connected in sequence in the liquid phase, wherein, The first gas-liquid separation device (210) is used to separate the product system to obtain a first gas component and hydrotreated oil; the first gas-liquid separation device (210) is provided with a first gas component outlet (211); The second gas-liquid separation device (220) is used to separate the hydrotreated oil to obtain a second gas component and a first oil product; the second gas-liquid separation device (220) is provided with a second gas component outlet (221); The first distillation unit (230) is selected from the first distillation column and is used to perform a first distillation on the first oil product to obtain the first naphtha, the first aviation kerosene, the diesel oil and the transformer oil; the side of the first distillation column is provided with the first naphtha outlet, the first aviation kerosene outlet (201), the diesel oil outlet (202) and the transformer oil outlet (203) from top to bottom; Preferably, the first separation unit (200) further includes a gas separation device (240), which is connected to the first gas component outlet (211) and is used to separate the first gas component to obtain hydrogen; more preferably, the gas separation device (240) is provided with a gas inlet and a hydrogen outlet, the gas inlet is connected to the first gas component outlet (211), and the hydrogen outlet is connected to at least one of the first hydrogen inlet (112), the second hydrogen inlet (121), and the third hydrogen inlet (131).

3. The coal direct liquefaction oil and bio-based oil co-hydrogenation treatment system according to claim 1 or 2, characterized in that, The second separation unit (400) includes an oil-water separation device (410) and a second distillation device (420) connected sequentially in the liquid phase, wherein, The oil-water separator (410) is provided with a water outlet (411). The oil-water separator (410) is used to separate the oil-water mixture to obtain water and a second oil product. The second distillation unit (420) is selected from the second distillation column and is used to perform a second distillation on the second oil product to obtain the second naphtha, the gasoline, the light aromatics and the second aviation kerosene; the side of the second distillation column is provided with the second naphtha outlet, the gasoline outlet (401), the light aromatics outlet (402) and the second aviation kerosene outlet (403) from top to bottom.

4. The coal direct liquefaction oil and bio-based oil co-hydrogenation treatment system according to any one of claims 1 to 3, characterized in that, The first catalyst is loaded at a volume percentage of 50-60 vol% in the hydrorefining unit (110); preferably, the first catalyst is selected from one or more of the group consisting of cobalt-manganese / alumina, nickel-tungsten / alumina, and nickel-molybdenum / alumina; and / or, The second catalyst is loaded at a volume percentage of 60-70 vol% of the hydroisomerization unit (120); preferably, the second catalyst is selected from one or more of the group consisting of platinum / ZSM-5 molecular sieve, platinum / SAPO-11 molecular sieve, and tungsten-molybdenum / β molecular sieve; and / or, The third catalyst is loaded at a volume percentage of 55-70 vol% of the hydrogenation saturation unit (130); preferably, the third catalyst is selected from one or more of the group consisting of palladium / alumina, platinum / alumina, and platinum-rhenium / alumina; and / or, The fourth catalyst is loaded at a volume percentage of 45-65 vol% of the co-hydrogenation treatment unit (300); preferably, the fourth catalyst is selected from one or more of the group consisting of nickel-molybdenum-tungsten / alumina, platinum-rhenium / alumina, and nickel-copper / alumina.

5. A method for co-hydrogenating coal direct liquefaction oil and bio-based oil, characterized in that, The co-hydrogenation treatment method is carried out using the coal direct liquefaction oil and bio-based oil co-hydrogenation treatment system as described in any one of claims 1 to 4. The co-hydrogenation treatment method includes: Step S1: The direct coal liquefaction oil and hydrogen are subjected to a hydrorefining reaction in the presence of a first catalyst using the hydrorefining device (110) in the hydrorefining unit (100) to obtain the hydrorefined product. The hydrogenation isomerization device (120) in the hydrogenation unit (100) is used to carry out a hydrogenation isomerization reaction with hydrogen under the catalysis of a second catalyst to obtain a hydrogenation isomer product. The hydrogenation isomer product and hydrogen are subjected to a hydrogenation saturation reaction under the catalysis of a third catalyst using the hydrogenation saturation device (130) in the hydrogenation unit (100) to obtain the product system. Step S2, the product system is separated by the first separation unit (200) to obtain first naphtha, first aviation kerosene with a distillation range of 160-270℃, diesel oil and transformer oil with a distillation range of 290-360℃; Step S3: Under the catalysis of the fourth catalyst, the first naphtha and the bio-based oil are subjected to a co-hydrogenation reaction in a co-hydrogenation treatment device (300) to obtain an oil-water mixture; In step S4, the oil-water mixture is separated using a second separation unit (400) to obtain second naphtha, gasoline, light aromatics, and second aviation kerosene with a distillation range of 170-240°C. The light aromatics include benzene, toluene, and xylene. The second naphtha is then returned to step S3 for recycling.

6. The method for co-hydrogenation of coal direct liquefaction oil and bio-based oil according to claim 5, characterized in that, In step S1, the hydrogenation refining reaction is carried out at a temperature of 280–360°C, a pressure of 4–8 MPa, and a volume hourly space velocity of 0.5–2.5 h⁻¹. -1 ; Preferably, the volume ratio of hydrogen to direct coal liquefaction oil is (1000-1500):1; Preferably, the hydroisomerization reaction is carried out at a temperature of 300–400°C, a pressure of 6–10 MPa, and a volume hourly space velocity of 0.5–1.5 h⁻¹. -1 ; Preferably, the volume ratio of hydrogen to the hydrorefined product is (500-700):1; Preferably, the hydrogenation saturation reaction is carried out at a temperature of 340–420°C, a pressure of 6–10 MPa, and a volume hourly space velocity of 0.5–3 h⁻¹. -1 ; Preferably, the volume ratio of hydrogen to the hydroisomer is (800-1000):

1.

7. The method for co-hydrogenation of coal direct liquefaction oil and bio-based oil according to claim 5 or 6, characterized in that, Step S2 includes: performing a first gas-liquid separation on the product system using a first gas-liquid separation device (210) to obtain a first gas component and hydrogenated oil; performing a second gas-liquid separation on the hydrogenated oil using a second gas-liquid separation device (220) to obtain a second gas component and a first oil product; and performing a first distillation on the first oil product using a first distillation device (230) to obtain the first naphtha, the first aviation kerosene, the diesel oil, and the transformer oil. Preferably, the temperature of the first gas-liquid separation is 30–60°C, the pressure is 6–10 MPa, and the time is 0.1–0.5 h; Preferably, the temperature of the second gas-liquid separation is 40–70°C, the pressure is 0.5–3 MPa, and the time is 0.1–0.5 h; Preferably, the first distillation is carried out using a first distillation column; more preferably, the top temperature of the first distillation column is 100-120°C, the top pressure is 0.1-0.3 MPa, the bottom temperature is 280-340°C, and the reflux ratio is (2-6):

1.

8. The method for co-hydrogenation of coal direct liquefaction oil and bio-based oil according to claim 7, characterized in that, Step S2 further includes separating the first gas component using a gas separation device (240) to obtain hydrogen; preferably, the hydrogen is recycled to at least one of the hydrogenation purification reaction, the hydrogenation isomerization reaction and the hydrogenation saturation reaction in step S1.

9. The method for co-hydrogenation treatment of coal direct liquefaction oil and bio-based oil according to any one of claims 5 to 8, characterized in that, In step S3, the co-hydrogenation reaction is carried out at a temperature of 300–420°C, a pressure of 0.5–1.5 MPa, and a volume hourly space velocity of 1–2.5 h⁻¹. -1 ; Preferably, the volume ratio of the first naphtha to the bio-based oil is (1.5-2):1; More preferably, the bio-based oil is selected from one or more of the group consisting of waste edible oil pyrolysis oil, microalgae pyrolysis oil, and lignin liquefaction oil.

10. The method for co-hydrogenation of coal direct liquefaction oil and bio-based oil according to claim 9, characterized in that, Step S4 includes: separating the oil-water mixture using an oil-water separator (410) to obtain water and a second oil product; and performing a second distillation on the second oil product using a second distillation apparatus (420) to obtain the second naphtha, the gasoline, the light aromatics, and the second aviation kerosene. Preferably, a second distillation column is used for the second distillation; more preferably, the top temperature of the second distillation column is 30-80°C, the top pressure is 0.1-0.5 MPa, the bottom temperature is 250-320°C, and the reflux ratio is (3-8):1.

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