Method for hydrogenating coking gasoline and diesel oil

Abstract

The present invention relates to the technical field of oil refining and discloses a method for hydrogenating coking gasoline and diesel. The method comprises the following steps: (1) subjecting the coking gasoline and diesel raw material oil to a first fractionation to obtain a raw material oil light fraction, a raw material oil middle fraction and a raw material oil heavy fraction; (2) carrying out a first hydrogenation reaction on the raw material oil heavy fraction and hydrogen in the presence of a first hydrogenation catalyst; (3) carrying out a second hydrogenation reaction on the first hydrogenation reaction product and the raw material oil middle fraction in the presence of a second hydrogenation catalyst; (4) carrying out a third hydrogenation reaction on the second hydrogenation reaction product and the raw material oil light fraction in the presence of a third hydrogenation catalyst; (5) carrying out gas-liquid separation on the third hydrogenation reaction product to obtain a gas phase and a liquid phase, and subjecting the liquid phase to stripping and a second fractionation to obtain a hydrogenated product. By controlling the temperature of the reaction zone, the method avoids excessive hydrodesulfurization.

Description

Technical Field

[0001] The present invention relates to the field of petroleum refining, and relates to a method for hydrogenating coker gasoline and diesel oil. Background Art

[0002] At present, higher requirements are relatively put forward for the secondary processing capacity of crude oil in petroleum refining enterprises. The secondary processing units of petroleum refining enterprises are mainly coking, fluid catalytic cracking, hydrogenation, catalytic reforming and other units. In particular, delayed coking has become an important means for heavy oil conversion.

[0003] Delayed coking is one of the important means for refining enterprises to process inferior heavy oil, but the products produced need to be further purified to remove impurities. In particular, coker gasoline contains high contents of impurities such as sulfur and olefins, and has poor stability and is not suitable as vehicle gasoline. However, coker naphtha can be used as a raw material for units such as ethylene cracking, fertilizer and reforming after hydrogenation treatment. Since coker naphtha itself contains high contents of impurities such as sulfur, nitrogen, olefins and gum, it must be first subjected to hydrofining to be used as a raw material for units such as ethylene cracking, fertilizer and reforming. And coker diesel also contains high contents of impurities such as sulfur, nitrogen and olefins. Coker diesel can be used to produce vehicle diesel after hydrofining according to its product characteristics. The article "Research on Hydrogenating Coker Gasoline and Diesel Oil to Produce Jet Fuel" on pages 173-176 of the 3rd issue of Petrochemical Technology & Application in 2004 introduced that coker gasoline and diesel oil were mixed for deep hydrofining, and then gasoline, kerosene and diesel oil were fractionated out. This method only deeply refined the whole coker fraction and then separated the kerosene fraction therein, resulting in over-hydrogenation. The article "Hydrofining Coker Gasoline and Diesel Oil under Low Pressure with FH-5 Catalyst" published by Sun Guang on pages 39-42 of the 5th issue of Petroleum Refinery Engineering in 1994 used a set of hydrofining devices to hydrofine coker gasoline and coker diesel. Specifically, a feed switching method was adopted, that is, after processing a batch of coker gasoline, the raw material was switched to coker diesel. This method can only be applied to small-scale coking units. During the production process, it is necessary to continuously switch, and the operating conditions need to be frequently changed, which brings inconvenience to the operation and has a great impact on the catalyst performance.

[0004] CN101003751A discloses a processing method for coking whole-fraction oil, including: separating coking whole-fraction oil into coking light-fraction oil and coking heavy-fraction oil, where the light-fraction oil contains part of the light diesel fraction and the heavy-fraction oil is a heavy diesel fraction with a higher final boiling point. The coking heavy-fraction oil is subjected to hydrocracking, and the cracking products are mixed with the coking light-fraction oil and then subjected to hydrofining treatment. CN111321005A discloses a hydroprocessing method for producing diesel with low energy consumption and long cycle. First, the diesel raw material is fractionated into a light fraction and a heavy fraction. The heavy fraction undergoes a hydrodesulfurization reaction in the first reaction zone and then reacts with the light-fraction diesel raw material oil in the second hydroprocessing reaction zone under the action of a hydrodesulfurization catalyst. CN200410050729.6 discloses a hydrotreating method for diesel fraction. First, the raw material oil is fractionated into a light fraction and a heavy fraction. The light fraction is sent to the first reaction zone and contacts with a hydrofining catalyst, and the heavy fraction is sent to the second reaction zone and contacts with a hydrofining catalyst. The liquid-phase products of the two reaction zones are mixed into a refined fraction oil. The raw material of this process needs to be fractionated into two groups of light and heavy components, and the process is relatively complex.

[0005] The processes selected in the above patents are all conventional gas-phase circulating trickle-bed hydroprocessing processes, but there are problems of excessive hydrodesulfurization, which easily causes over-quality of products, reduction of liquid yield, excessive hydrogen consumption and energy consumption, etc. Summary of the Invention

[0006] The object of the present invention is to overcome the problems of excessive hydrodesulfurization and the like existing in the prior art, and provide a method for hydroprocessing coking gasoline and diesel. This method can better couple the reaction fractions, reaction sequence and reaction temperature in depth, that is, realize reactions such as desulfurization, desiliconization and deolefination, and desulfurization of mercaptan, and avoid excessive hydrodesulfurization reaction, reduce hydrogen consumption, and increase liquid yield; at the same time, reduce the reaction severity, and can better play the role of each reaction zone, make better use of the heat balance, reduce the energy consumption of the device, and produce high-quality diesel, jet fuel and naphtha for ethylene cracking, and is particularly suitable for the mixed hydroprocessing of coking gasoline and diesel.

[0007] To achieve the above object, the present invention provides a method for hydroprocessing coking gasoline and diesel, which includes the following steps:

[0008] (1) Subjecting the coking gasoline and diesel raw material oil to the first fractionation to obtain a raw material oil light fraction, a raw material oil middle fraction and a raw material oil heavy fraction;

[0009] (2) In the presence of a first hydroprocessing catalyst, subjecting the raw material oil heavy fraction obtained in step (1) to a first hydroprocessing reaction with hydrogen to obtain a first hydroprocessing reaction product;

[0010] (3) In the presence of a second hydroprocessing catalyst, subjecting the first hydroprocessing reaction product obtained in step (2) and the raw material oil middle fraction obtained in step (1) to a second hydroprocessing reaction to obtain a second hydroprocessing reaction product;

[0011] (4) In the presence of a third hydrogenation catalyst, the second hydrogenation reaction product obtained in step (3) and the light fraction of the feedstock oil obtained in step (1) are subjected to a third hydrogenation reaction to obtain a third hydrogenation reaction product;

[0012] (5) The third hydrogenation reaction product obtained in step (4) is subjected to gas-liquid separation to obtain a gas phase and a liquid phase, and the liquid phase is subjected to stripping and second fractionation to obtain a hydrogenation product;

[0013] Wherein, in step (2), the average reaction temperature of the first hydrogenation reaction is 340 - 420 °C;

[0014] In step (3), the average reaction temperature of the second hydrogenation reaction is 260 - 310 °C;

[0015] In step (4), the average reaction temperature of the third hydrogenation reaction is 240 - 300 °C.

[0016] Preferably, in step (2), the conditions of the first hydrogenation reaction include: hydrogen partial pressure is 3 - 16 MPa, preferably 4 - 16 MPa, volume space velocity is 0.3 - 10 h 1 , and the hydrogen-oil volume ratio is 200:1 to 2500:1.

[0017] Preferably, in step (2), the organic sulfur content in the reaction effluent of the first hydrogenation reaction is 5 - 10 μg / g.

[0018] Preferably, in step (2), the hydrogen sulfide concentration in the reaction effluent of the first hydrogenation reaction is 4000 - 14000 ppm, preferably 5000 - 10000 ppm.

[0019] Preferably, in step (2), the first hydrogenation catalyst includes a first carrier and a first hydrogenation active component supported on the first carrier, and the first hydrogenation active component is W-Ni, Mo-Ni or W-Mo-Ni.

[0020] Preferably, in step (2), the first carrier is alumina, silica, amorphous silica-alumina, titanium oxide, or their composite oxides or mixtures.

[0021] Preferably, in step (2), based on the weight of the first hydrogenation catalyst, the content of the first hydrogenation active component in terms of oxide is 15 - 50 wt%, preferably 18 - 45 wt%.

[0022] Preferably, in step (2), the specific surface area of the first hydrogenation catalyst is 120 - 600 m 2 / g, preferably 200 - 450 m 2 / g, the pore volume is 0.2 - 1.5 mL / g, preferably 0.35 - 0.7 mL / g.

[0023] Preferably, in step (3), the conditions of the second hydrogenation reaction include: the hydrogen partial pressure is 1 - 10 MPa, preferably 2 - 8 MPa, the volume space velocity is 0.6 - 10 h 1 , and the hydrogen-oil volume ratio is 100:1 - 2500:1.

[0024] Preferably, in step (3), the organic sulfur content in the reaction effluent of the second hydrogenation reaction is 50 - 2000 μg / g.

[0025] Preferably, in step (3), the second hydrogenation catalyst includes a second carrier and a second hydrogenation active component supported on the second carrier, and the second hydrogenation active component is a metal oxide of Group VIB or / and a metal oxide of Group VIII.

[0026] Preferably, in step (3), based on the weight of the second hydrogenation catalyst, the content of the metal oxide of Group VIB is 5 - 30 wt%, preferably 5 - 15 wt%, and the content of the metal oxide of Group VIII is 1 - 15 wt%, preferably 1.5 - 6 wt%.

[0027] Preferably, in step (3), the specific surface area of the second hydrogenation catalyst is 100 - 550 m 2 / g, preferably 150 - 420 m 2 / g, the pore volume is 0.3 - 1.5 mL / g, preferably 0.4 - 0.9 mL / g.

[0028] Preferably, in step (4), the conditions of the third hydrogenation reaction include: the hydrogen partial pressure is 0.5 - 8 MPa, preferably 1.5 - 6 MPa, the volume space velocity is 0.6 - 10 h 1 , and the hydrogen-oil volume ratio is 100:1 - 2500:1.

[0029] Preferably, in step (4), the organic sulfur content in the reaction effluent of the third hydrogenation reaction is 50 - 2000 μg / g.

[0030] Preferably, in step (4), the third hydrogenation catalyst includes a third carrier and a third hydrogenation active component supported on the third carrier, and the third hydrogenation active component is a metal oxide of Group VIB or / and a metal oxide of Group VIII, preferably the metal oxide of Group VIII is Co.

[0031] Preferably, in step (4), based on the weight of the third hydrogenation catalyst, the content of the Group VIB metal oxide is 5-30wt%, preferably 5-15wt%, and the content of the Group VIII metal oxide is 1-15wt%, preferably 1.5-6wt%.

[0032] Preferably, in step (4), the specific surface area of the third hydrogenation catalyst is 80-500m 2 / g, preferably 100-380m 2 / g, and the pore volume is 0.3-1.5 mL / g, preferably 0.4-0.9 mL / g.

[0033] Preferably, the average reaction temperature of the first hydrogenation reaction is 340-400°C; the average reaction temperature of the second hydrogenation reaction is 260-310°C; and the average reaction temperature of the third hydrogenation reaction is 240-300°C.

[0034] Preferably, the average reaction temperature of the second hydrogenation reaction is 40-150°C lower than that of the first hydrogenation reaction, preferably 50-100°C lower, and the average reaction temperature of the third hydrogenation reaction is 0-70°C lower than that of the second hydrogenation reaction, preferably 5-50°C lower.

[0035] Preferably, the hydrogen partial pressure of the second hydrogenation reaction is 2-6 MPa lower than that of the first hydrogenation reaction, and the hydrogen partial pressure of the third hydrogenation reaction is 1-4 MPa lower than that of the second hydrogenation reaction.

[0036] Preferably, the first hydrogenation reaction, the second hydrogenation reaction and the third hydrogenation reaction are carried out in the same reactor.

[0037] Preferably, the volume ratio of the first hydrogenation catalyst, the second hydrogenation catalyst and the third hydrogenation catalyst is (40-80):(10-40):(10-20), preferably (75-50):(15-25):(10-25).

[0038] Preferably, in step (1), the initial distillation point temperature of the coking gasoline and diesel feedstock oil is 30-100°C, and the final distillation point temperature is 350-400°C.

[0039] Preferably, the cut points of the feedstock oil light fraction, the feedstock oil middle fraction and the feedstock oil heavy fraction are 150-200° C. and 220-270° C. respectively.

[0040] Preferably, the method further comprises: removing hydrogen sulfide from the gas phase obtained in step (5), returning the gas phase to step (1) and mixing it with coking gasoline and diesel feedstock oil.

[0041] Preferably, the hydrogen sulfide concentration in the material obtained in step (5) after hydrogen sulfide is removed from the gas phase is 0-1500 ppm, preferably 50-1000 ppm.

[0042] Preferably, the hydrogenation products are gasoline, diesel and jet fuel.

[0043] Preferably, the sulfur content in the gasoline is 25 - 1500 ppm; the sulfur content in the jet fuel is 500 - 1900 ppm.

[0044] Through the above technical solution, the method for hydrogenating coking gasoline and diesel of the present invention has the following beneficial effects:

[0045] 1. The coking gasoline and diesel processing technology provided by the present invention divides the feedstock oil into different fractions, optimizes the catalyst grading according to the different reaction requirements of the reaction materials, deeply combines the reaction fractions, reaction sequence, hydrogen partial pressure, hydrogen sulfide concentration and various reactions, achieves precise control of the reaction depth, balances the reaction heat, gives full play to the performance of the catalyst, that is, completes the deep desulfurization of the diesel fraction, and controls the strong exothermic reaction of olefins to avoid excessive reaction temperature rise, and avoids the passive desulfurization of the jet fuel fraction and gasoline fraction, so as to achieve the purpose of progressive controllability of the reaction process.

[0046] 2. In the hydrogenation method provided by the present invention, the heavy fraction first enters the first reaction zone for hydrofining desulfurization reaction alone, avoiding the strong exothermic reaction of a large amount of olefins in the light fraction, reducing the reaction temperature rise, and avoiding the situation of too high peak reaction temperature and the passive desulfurization reaction of the gasoline fraction and jet fuel fraction. At the same time, the product is used as a heat-carrying medium and mixed with the gasoline fraction and jet fuel fraction to enter the second and third reaction zones, increasing the heat capacity and controlling the strong exothermic reaction of olefins in the gasoline fraction to cause the passive increase of the reaction temperature rise. The reaction product is mixed with the gasoline fraction and the jet fuel fraction, which can further reduce the reaction temperature, realize the de-olefination, de-mercaptan and partial desulfurization reactions at a low reaction temperature. This reaction process is a gentle reaction, avoiding the passive over-hydrodesulfurization when the light components and heavy components are mixed and reacted, and reducing the problems of reduction of the target product yield and increase of hydrogen consumption caused by over-reaction.

[0047] 3. In the hydrogenation method provided by the present invention, hydrofining catalysts, hydro-de-olefination catalysts and low-temperature hydro-desulfurization mercaptan catalysts are respectively installed in the first, second and third reaction zones, and the corresponding reaction temperatures and hydrogen partial pressures are adjusted. Following the rule that the reactants are from difficult to easy, and the reaction temperature and hydrogen partial pressure are from high to low, it gives full play to the high-temperature hydrodesulfurization, low-temperature hydro-desulfurization mercaptan activity and de-olefination performance of the catalyst, and avoids the passive over-desulfurization of sulfur in the gasoline and jet fuel fractions in the conventional coking gasoline and diesel hydrogenation process to ensure ultra-deep removal of thiophene sulfur. This can reduce the sulfur injection amount of the downstream ethylene plant, improve the total liquid yield while increasing the yields of gasoline and jet fuel components, and reduce hydrogen consumption.

[0048] 4. In the hydrogenation method provided by the present invention, the heavy fraction enters the first reaction zone and undergoes a deep desulfurization reaction with high-purity hydrogen under the action of a catalyst, reducing the inhibition of the reaction caused by excessive hydrogen sulfide concentration, and it is easier to obtain high-quality diesel components. The hydrogen sulfide concentration in the product of the first reaction zone is then mixed with the light fraction for hydrogenation, which can achieve the purpose of inhibiting the depth of desulfurization, and avoid the exotherm of olefins in the light fraction, resulting in an increase in the depth of desulfurization, reducing the sulfur removal rate of gasoline and jet fuel fractions that are too high, reducing the sulfur injection amount of the downstream ethylene plant, and increasing the yield of jet fuel components.

[0049] 5. The hydrogenation method provided by the present invention gradually controls the reaction exotherm, effectively matches the heat capacity and reaction heat balance, effectively controls the reaction temperature rise, reduces the probability of thermal cracking, and ensures high liquid yield, low hydrogen consumption and energy consumption.

[0050] 6. The method of the present invention can be slightly modified using existing coking gasoline and diesel hydrogenation or other similar hydrogenation devices, reducing the transformation cost. At the same time, the operation steps change little, the process is simple, the operation is convenient, the safety is good, the environment is friendly, and it has a long operation cycle. BRIEF DESCRIPTION OF THE DRAWINGS

[0051] Figure 1 It is a schematic diagram of the process adopted in the embodiment of the present invention.

[0052] DESCRIPTION OF THE REFERENCE NUMERALS

[0053] 1. Heavy fraction introduction pipeline 2. Fresh hydrogen introduction pipeline 3. Second hydrogen circulation pipeline

[0054] 5. Hydrofining reactor 6. Hydrogenation product export pipeline 7. Cold high-pressure separator

[0055] 8. Liquid phase pipeline 9. Cold low-pressure separator 10. Export pipeline

[0056] 11. First hydrogen circulation pipeline 13. Intermediate fraction introduction pipeline 14. Light fraction introduction pipeline DETAILED DESCRIPTION OF THE INVENTION

[0057] In the ranges disclosed herein, the endpoints and any values are not limited to the exact ranges or values, and these ranges or values should be understood to include values close to these ranges or values. For numerical ranges, between the endpoint values of each range, between the endpoint values of each range and individual point values, and between individual point values can be combined with each other to obtain one or more new numerical ranges, and these numerical ranges should be regarded as specifically disclosed herein.

[0058] The method for coking gasoline and diesel hydrogenation provided by the present invention includes the following steps:

[0059] (1) The coking gasoline and diesel raw material oil is fractionated for the first time to obtain a raw material oil light fraction, a raw material oil middle fraction, and a raw material oil heavy fraction;

[0060] (2) In the presence of a first hydrogenation catalyst, the raw material oil heavy fraction obtained in step (1) reacts with hydrogen in a first hydrogenation reaction to obtain a first hydrogenation reaction product;

[0061] (3) In the presence of a second hydrogenation catalyst, the first hydrogenation reaction product obtained in step (2) and the raw material oil middle fraction obtained in step (1) undergo a second hydrogenation reaction to obtain a second hydrogenation reaction product;

[0062] (4) In the presence of a third hydrogenation catalyst, the second hydrogenation reaction product obtained in step (3) and the raw material oil light fraction obtained in step (1) undergo a third hydrogenation reaction to obtain a third hydrogenation reaction product;

[0063] (5) The third hydrogenation reaction product obtained in step (4) is subjected to gas-liquid separation to obtain a gas phase and a liquid phase, and the liquid phase is subjected to stripping and second fractionation to obtain a hydrogenated product;

[0064] Among them, in step (2), the average reaction temperature of the first hydrogenation reaction is 340 - 420 °C;

[0065] In step (3), the average reaction temperature of the second hydrogenation reaction is 260 - 310 °C;

[0066] In step (4), the average reaction temperature of the third hydrogenation reaction is 240 - 300 °C.

[0067] In the present invention, the first hydrogenation reaction in step (2) is mainly used for hydrodesulfurization, the second hydrogenation reaction in step (3) is mainly used for hydrodealkylation reaction and mild desulfurization reaction, and the third hydrogenation reaction in step (4) is mainly used for hydrodesulfurization of mercaptan and desulfurization reaction.

[0068] In the present invention, the coking gasoline and diesel raw material oil can be one or more of various hydrocarbons containing gasoline and / or diesel fractions, specifically a mixture of more than one of C4 - C25 hydrocarbons. In addition, the raw material oil can also be mixed with and processed one or more of straight-run diesel, straight-run jet fuel, catalytic gasoline, catalytic diesel, ethylene cracking gasoline, straight-run naphtha and other oil products.

[0069] In the present invention, for the coking gasoline and diesel raw material oil, the coking gasoline and diesel raw material oil commonly used for hydrogenation can be used. The initial boiling point temperature of the coking gasoline and diesel raw material oil can be 30 - 100 °C, preferably 40 - 90 °C, more preferably 60 - 90 °C, and the final boiling point temperature can be 340 - 400 °C, preferably 350 - 380 °C.

[0070] According to the present invention, in step (1), the coking gasoline and diesel raw material oil is separated into a raw material oil light fraction, a raw material oil middle fraction, and a raw material oil heavy fraction by the first fractionation. As the cut points for the raw material oil light fraction, the raw material oil middle fraction, and the raw material oil heavy fraction, 150 - 200 °C and 220 - 270 °C are preferred, and 165 - 190 °C and 230 - 250 °C are more preferred. By performing fractionation under the above conditions, the yield of the obtained gasoline can be increased, and the economic benefits can be improved.

[0071] According to the present invention, in step (2), the conditions of the first hydrogenation reaction may include: the hydrogen partial pressure is 3 - 16 MPa, preferably 4 - 16 MPa, the volume space velocity is 0.3 - 10 h 1 , preferably 0.6 - 2 h 1 , and the hydrogen-oil volume ratio is 200:1 to 2500:1, preferably 300:1 to 1000:1.

[0072] According to the present invention, preferably, the organic sulfur content in the reaction effluent in the first reaction zone is 5 - 10 μg / g.

[0073] According to the present invention, preferably, the hydrogen sulfide concentration in the reaction effluent of the first hydrogenation reaction is 4000 - 14000 ppm, preferably 5000 - 10000 ppm.

[0074] According to the present invention, in step (2), the first hydrogenation catalyst can be selected as a suitable commercial catalyst according to the needs of the process flow, can also be prepared by a conventional method, or can be a regenerated catalyst after the deactivated catalyst is regenerated. The first hydrogenation catalyst includes a first carrier and a first hydrogenation active component supported on the first carrier. The first hydrogenation active component can be W-Ni, Mo-Ni, or W-Mo-Ni. The first carrier is generally a refractory porous oxide, such as alumina, silica, amorphous silica-alumina, titanium oxide, or their composite oxide or mixed oxide carriers, etc. Among them, a material without acidity or weak acidity is generally used as the carrier. Preferably, based on the weight of the first hydrogenation catalyst, the content of the first hydrogenation active component in terms of oxide can be 15 - 50 wt%, preferably 18 - 45 wt%, and the nickel oxide content is 2 - 8 wt%, preferably 2.3 - 6 wt%.

[0075] The shape of the first hydrogenation catalyst is not particularly limited. For example, it can be spherical or bar-shaped. The spherical diameter is 0.04 - 12 mm, preferably 0.04 - 5 mm; the bar shape has a length of 2 - 15 mm, preferably 2 - 8 mm, and a diameter of 1 - 6 mm, preferably 1.5 - 3.5 mm.

[0076] Preferably, the specific surface area of the first hydrogenation catalyst is 120 - 600 m 2 / g, preferably 200 - 450 m 2 / g, the pore volume is 0.2 - 1.5 mL / g, preferably 0.35 - 0.7 mL / g.

[0077] According to the present invention, in step (3), the conditions for the second hydrogenation reaction may include: the hydrogen partial pressure is 1 - 10 MPa, preferably 2 - 8 MPa, the volume space velocity is 0.6 - 10 h 1 , and the hydrogen - to - oil volume ratio is 100:1 - 2500:1, preferably 200:1 - 1000:1.

[0078] According to the present invention, preferably, the organic sulfur content in the reaction effluent of the second hydrogenation reaction is 50 - 2000 μg / g.

[0079] According to the present invention, in step (3), the second hydrogenation catalyst can be selected as a suitable commercial catalyst according to the needs of the process flow, can also be prepared by existing methods, or can be a regenerated catalyst after the deactivated catalyst is regenerated. The second hydrogenation catalyst includes a second carrier and a second hydrogenation active component supported on the second carrier. The second hydrogenation active component is a metal oxide of Group VIB or / and a metal oxide of Group VIII. For example, it can be Mo, W, Co, Ni, etc. The second carrier can be alumina or alumina modified with additives. Preferably, based on the weight of the second hydrogenation catalyst, the content of the metal oxide of Group VIB can be 5 - 30 wt%, preferably 5 - 15 wt%, and the content of the metal oxide of Group VIII can be 1 - 15 wt%, preferably 1.5 - 6 wt%.

[0080] The shape of the second hydrogenation catalyst is not particularly limited. For example, it can be spherical or bar - shaped. The spherical diameter is 0.2 - 20 mm, preferably 0.3 - 5 mm; the bar - shaped one has a length of 2 - 15 mm, preferably 3 - 8 mm, and a diameter of 1.5 - 6 mm, preferably 2.0 - 5.5 mm.

[0081] Preferably, the specific surface area of the second hydrogenation catalyst is 100 - 550 m 2 / g, preferably 150 - 420 m 2 / g, the pore volume is 0.3 - 1.5 mL / g, preferably 0.4 - 0.9 mL / g.

[0082] According to the present invention, in step (4), the conditions for the third hydrogenation reaction include: the hydrogen partial pressure is 0.5 - 8 MPa, preferably 1.5 - 6 MPa, the volume space velocity is 0.6 - 10 h 1 , and the hydrogen - to - oil volume ratio is 100:1 - 2500:1, preferably 150:1 - 1000:1.

[0083] According to the present invention, in step (4), the third hydrogenation catalyst can be a suitable commercial catalyst selected according to the needs of the process flow, can also be prepared by existing methods, or can be a regenerated catalyst after the deactivated catalyst is regenerated. The third hydrogenation catalyst includes a third carrier and a third hydrogenation active component supported on the third carrier. The third hydrogenation active component is a metal oxide of Group VIB or / and a metal oxide of Group VIII. For example, it can be Mo, W, Co, etc. Preferably, the metal oxide of Group VIII is Co. The third carrier can be alumina or alumina modified with an auxiliary agent. Preferably, based on the weight of the third hydrogenation catalyst, the content of the metal oxide of Group VIB can be 5-30 wt%, preferably 5-15 wt%, and the content of the metal oxide of Group VIII can be 1-15 wt%, preferably 1.5-6 wt%.

[0084] The shape of the third hydrogenation catalyst is not particularly limited. For example, it can be spherical or strip-shaped. The spherical diameter is 0.2-20 mm, preferably 0.3-5 mm; the strip shape has a length of 2-15 mm, preferably 3-8 mm, and a diameter of 1.5-6 mm, preferably 2.0-5.5 mm.

[0085] Preferably, the specific surface area of the third hydrogenation catalyst is 80-500 m 2 / g, preferably 100-380 m 2 / g, and the pore volume is 0.3-1.5 mL / g, preferably 0.4-0.9 mL / g.

[0086] According to some preferred embodiments of the present invention, from the perspective of facilitating the hydrode-olefination reaction and mild desulfurization reaction in the second hydrogenation reaction of step (3) and the hydrodesulfurization of mercaptan and desulfurization reaction in the third hydrogenation reaction of step (4), preferably, the second hydrogenation catalyst is a molybdenum-nickel-based catalyst, and the third hydrogenation catalyst is a molybdenum-cobalt-based catalyst.

[0087] According to the present invention, the preparation methods of the first hydrogenation catalyst, the second hydrogenation catalyst, and the third hydrogenation catalyst are well-known in the art. One or more of impregnation method, co-extrusion method, and co-precipitation method can be used, or it can also be a regenerated catalyst after the deactivated catalyst is regenerated. In order to obtain a better hydrogenation catalytic effect, preferably, the volume ratio of the first hydrogenation catalyst, the second hydrogenation catalyst, and the third hydrogenation catalyst is (40-80):(10-40):(10-20), preferably (75-50):(15-25):(10-25), and more preferably (60-50):(20-30):(20-25).

[0088] In the present invention, the content of organic sulfur in the reaction effluent in the third reaction zone is within the above range, which can avoid excessive hydrogenation of coking gasoline and diesel feedstock, thus being more conducive to the application of the product as a feedstock for ethylene cracking. Downstream ethylene cracking units need to inject sulfur-containing media (such as dimethyl sulfide (DMDS)) when using feedstocks such as hydrotreated naphtha, hydrocracked light naphtha, hydrocracked oil, and ethane. If the ethylene cracking feedstock contains a sufficient amount of sulfur, sulfur injection is not required. The ethylene cracking unit needs to maintain the sulfur concentration in the feedstock above 100 ppm, and the sulfur injection amount is calculated based on the feedstock flow rate of the cracking furnace.

[0089] According to some more preferred embodiments of the present invention, the average reaction temperature of the first hydrogenation reaction can be 420 °C or lower, 410 °C or lower, 400 °C or lower, 390 °C or lower, 380 °C or lower, 370 °C or lower, 360 °C or lower, or 350 °C or lower, and 340 °C or higher, 350 °C or higher, 360 °C or higher, 370 °C or higher, 380 °C or higher, 390 °C or higher, or 400 °C or higher, preferably 340 - 370 °C; the average reaction temperature of the second hydrogenation reaction is 310 °C or lower, 300 °C or lower, 290 °C or lower, 280 °C or lower, or 270 °C or lower, and 260 °C or higher, 270 °C or higher, 280 °C or higher, 290 °C or higher, or 300 °C or higher, preferably 260 - 310 °C; the average reaction temperature of the third hydrogenation reaction is 300 °C or lower, 290 °C or lower, 280 °C or lower, 270 °C or lower, 260 °C or lower, or 250 °C or lower, and 240 °C or higher, 250 °C or higher, 260 °C or higher, 270 °C or higher, 280 °C or higher, or 290 °C or higher, preferably 240 - 300 °C. Further preferably, the average reaction temperature of the second hydrogenation reaction is 40 - 150 °C lower than that of the first hydrogenation reaction, for example, it can be 50 - 100 °C or 60 - 80 °C lower, and the average reaction temperature of the third hydrogenation reaction is 0 - 70 °C lower than that of the second hydrogenation reaction, preferably 5 - 50 °C lower, for example, it can be 0 - 20 °C, 5 - 20 °C, or 5 - 15 °C lower, etc.

[0090] According to the present invention, the hydrogen partial pressure of the first hydrogenation reaction can be, for example, 3 MPa, 4 MPa, 5 MPa, 6 MPa, 7 MPa, 8 MPa, 9 MPa, 10 MPa, 11 MPa, 12 MPa, 13 MPa, 14 MPa, 15 MPa, or 16 MPa; the hydrogen partial pressure of the second hydrogenation reaction can be, for example, 2 MPa, 3 MPa, 4 MPa, 5 MPa, 6 MPa, 7 MPa, 8 MPa, 9 MPa, or 10 MPa; the hydrogen partial pressure of the third hydrogenation reaction can be 1 MPa, 2 MPa, 3 MPa, 4 MPa, 5 MPa, 6 MPa, 7 MPa, 8 MPa, 9 MPa, or 10 MPa.

[0091] Preferably, the hydrogen partial pressure of the second hydrogenation reaction is 2-6 MPa lower than that of the first hydrogenation reaction. For example, it can be 2 MPa, 2.5 MPa, 3 MPa, 3.5 MPa, 4 MPa, 4.5 MPa, 5 MPa, 5.5 MPa or 6 MPa lower. The hydrogen partial pressure of the third hydrogenation reaction is 0-4 MPa lower than that of the second hydrogenation reaction. For example, it can be 0 MPa, 0.5 MPa, 1 MPa, 1.5 MPa, 2 MPa, 2.5 MPa, 3 MPa, 3.5 MPa or 4 MPa lower.

[0092] By adopting the above reaction conditions, the sulfur content of the obtained hydrogenation product can be further regulated.

[0093] As the device for carrying out the above first hydrogenation reaction, second hydrogenation reaction and third hydrogenation reaction, any device used for hydrogenation of coking gasoline and diesel raw oil in the art can be adopted, and a trickle bed reactor is preferably adopted. In the above method, the first reaction zone for the first hydrogenation reaction, the second reaction zone for the second hydrogenation reaction, or the third reaction zone for the third hydrogenation reaction can each independently be a reactor or a part of a reactor, or can be composed of multiple reactors. If multiple reactors are selected, preferably 2-4 reactors. Inside the first reaction zone, the second reaction zone or the third reaction zone, one catalyst bed layer can be independently arranged, or multiple catalyst bed layers can be arranged. If multiple catalyst bed layers are selected, preferably 2-5 catalyst bed layers.

[0094] According to the present invention, the first hydrogenation reaction, the second hydrogenation reaction and the third hydrogenation reaction are carried out in the same reactor. That is, the first hydrogenation reaction is carried out in the first reaction zone, the second hydrogenation reaction is carried out in the second reaction zone, the third hydrogenation reaction is carried out in the third reaction zone, and the first reaction zone, the second reaction zone and the third reaction zone are arranged in the same reactor. For example, the process shown in Figure 1 can be adopted, Figure 1 in which equipment such as oil pumps, valves, heating furnaces, etc. are omitted. The reaction device includes a hydrofining reactor 5, which successively includes a first reaction zone, a second reaction zone and a third reaction zone from top to bottom. The first, second and third hydrogenation catalysts are respectively arranged in the three reaction zones. At this time, when the catalyst is fully filled, the volume ratio of the first hydrogenation catalyst, the second hydrogenation catalyst and the third hydrogenation catalyst can be represented by the volume ratio of the first reaction zone, the second reaction zone and the first reaction zone.

[0095] According to the present invention, the hydrogen in the first hydrogenation reaction, the second hydrogenation reaction and the third hydrogenation reaction can come from the hydrogenation reaction product of the previous step, or can be respectively introduced into the corresponding reaction zones.

[0096] According to the present invention, the method may further include: after the hydrogen sulfide in the gas phase obtained in step (5) is removed, it is returned to step (1) and mixed with the coking gasoline and diesel raw material oil. Specifically, the alcohol amine adsorption method can be used for desulfurization. Preferably, the hydrogen sulfide concentration in the gas phase obtained after the hydrogen sulfide is removed in step (5) (i.e., the recycle hydrogen at the inlet of the first reaction zone) is 1500 ppm or less, preferably 50 - 1000 ppm.

[0097] According to the present invention, the hydrogenation product obtained by the hydrogenation method of the present invention can be gasoline, diesel, and jet fuel. In step (5), the fractionation system used in the second fractionation is a system that can accurately fractionate gasoline, jet fuel, and diesel.

[0098] By hydrogenating the coking gasoline and diesel raw material oil using the hydrogenation method of the present invention, diesel with a sulfur content of 10 ppm or less (preferably 2 - 8 ppm), jet fuel with a sulfur content of 400 - 1900 ppm (preferably 500 - 1800 ppm, more preferably 500 - 1000 ppm), and gasoline with a sulfur content of 25 - 1500 ppm (preferably 50 - 1500 ppm, more preferably 100 - 500 ppm, further preferably 100 - 500 ppm) can be obtained. In addition, the olefin content of the gasoline can reach 1% by volume or less, and the mercaptan content in the jet fuel is 10 ppm or less, preferably 2 - 6 ppm. Thus, it can be seen that the hydrogenation method of the present invention can avoid over-hydrogenation of the coking gasoline and diesel raw material oil, which is more conducive to the application of the product as an ethylene cracking raw material.

[0099] According to a preferred embodiment of the present invention, the method further includes returning the gas phase obtained in step (5) as recycle hydrogen to step (2), step (3), and step (4). Specifically, the gas phase obtained in step (5) can be desulfurized by a desulfurization tower and then returned to step (2), step (3), and step (4).

[0100] According to an embodiment of the present invention, when hydrogenating coking gasoline and diesel, such as Figure 1As shown in the figure, the following process can be adopted: The heavy fraction of coker gasoline and diesel is introduced through pipeline 1, mixed with the fresh hydrogen in the fresh hydrogen introduction pipeline 2 and the recycle hydrogen in the second hydrogen recycle pipeline 3, and enters from the top of the hydrofining reactor 5. The material passes through the first reaction zone and undergoes the first hydrogenation reaction in the presence of the first hydrogenation catalyst. The light fraction of coker gasoline and diesel is introduced through pipeline 4, mixed with the recycle hydrogen in the hydrogen recycle pipeline 3, and enters the middle of the hydrofining reactor 5, where it is mixed with the reaction effluent from the upper first reaction zone and enters the second reaction zone. In the presence of the second hydrogenation catalyst, hydrodesulfurization and hydrodearomatization are carried out. The mixed material flows out from the bottom of the hydrofining reactor 5; it enters the cold high-pressure separator 7 through the hydrogenation product export pipeline 6. The obtained liquid phase passes through the liquid phase pipeline 8, enters the cold low-pressure separator 9, and then enters the fractionation unit through the export pipeline 10; the recycle hydrogen separated by the cold high-pressure separator 7 enters the recycle hydrogen desulfurization system (including a desulfurization tower) through the first hydrogen recycle pipeline 11, and then returns to the hydrogen recycle pipeline 3 through the recycle hydrogen compressor 12.

[0101] The present invention will be described in detail below through examples. In the following examples, FHUDS-8 is the FHUDS-8 hydrofining catalyst (the active component is Mo-Ni) produced by the Fushun Branch of Sinopec Catalyst Company, FH-40C is the FH-40C hydrodearomatization catalyst (the active component is W-Mo-Ni) produced by the Fushun Branch of Sinopec Catalyst Company, and FH-40B is the FH-40B low-temperature mercaptan hydrodesulfurization catalyst (the active component is Mo-Co) produced by the Fushun Branch of Sinopec Catalyst Company.

[0102] Examples 1-8 and Comparative Examples 1-3

[0103] Adopt the device as Figure 1 shown to carry out the hydrogenation of coker gasoline and diesel. The hydrofining reactor 5 is a trickle-bed reactor, which successively includes a first reaction zone, a second reaction zone, and a third reaction zone from top to bottom. The hydrogenation catalysts in Table 2-4 are respectively arranged in the three reaction zones.

[0104] The coking gasoline and diesel raw oil shown in Table 1 is fractionated at the splitting points of 180 °C and 250 °C respectively to obtain raw oil light fraction, raw oil middle fraction and raw oil heavy fraction. The hydrogenation reaction is carried out using the hydrofining reactor 5. The coking gasoline and diesel heavy fraction is mixed with the fresh hydrogen in the fresh hydrogen introduction pipeline 2 and the recycle hydrogen in the second hydrogen recycle pipeline 3 through the heavy fraction introduction pipeline 1, and enters from the top of the hydrofining reactor 5. The material passes through the first reaction zone and undergoes the first hydrogenation reaction in the presence of the first hydrogenation catalyst. The coking gasoline and diesel middle fraction is mixed with the recycle hydrogen in the hydrogen recycle pipeline 3 through the middle fraction introduction pipeline 13 and enters the hydrofining reactor 5 between the first reaction zone and the second reaction zone, mixes with the reaction effluent of the first reaction zone, and enters the second reaction zone, and undergoes the second hydrogenation reaction in the presence of the second hydrogenation catalyst. The coking gasoline and diesel light fraction is mixed with the recycle hydrogen in the hydrogen recycle pipeline 3 through the light fraction introduction pipeline 14 and enters the lower part of the hydrofining reactor 5 between the second reaction zone and the third reaction zone, mixes with the reaction effluent of the second reaction zone, and enters the third reaction zone, and undergoes the third hydrogenation reaction in the presence of the third hydrogenation catalyst. The mixed material flows out from the bottom of the hydrofining reactor 5; it enters the cold high-pressure separator 7 through the hydrogenation product export pipeline 6, the obtained liquid phase passes through the liquid phase pipeline 8, enters the cold low-pressure separator 9, and then enters the fractionation unit through the export pipeline 10; the recycle hydrogen separated by the cold high-pressure separator 7 enters the recycle hydrogen desulfurization system through the first hydrogen recycle pipeline 11, is desulfurized using the desulfurization tower, and then returns to the hydrogen recycle pipeline 3 through the recycle hydrogen compressor 12.

[0105] The raw oil used in the examples and comparative examples of the present invention is shown in Table 1. The main operating process conditions and product properties in the examples and comparative examples of the present invention are shown in Tables 2-4.

[0106] Table 1 Properties of Raw Oil

[0107] Coked gasoline and diesel <![CDATA[Density (20 °C), g / cm 3 > 0.8351 Distillation range, °C (ASTM D86) Initial boiling point / 10% 45 / 95 50％ / 90％ 256 / 355 95% / Dry point (final boiling point) 369 / 375 Sulfur, μg / g 10655 Nitrogen, μg / g 816 Olefins, v% 19

[0108] Table 2

[0109] Item Example 1 Example 2 Example 3 Example 4 First reaction zone FHUDS-8 FHUDS-8 FHUDS-8 FHUDS-8 Second reaction zone FH-40C FH-40C FH-40C FH-40C Third reaction zone FH-40B FH-40B FH-40B FH-40B Hydrogen partial pressure in the first / second / third reaction zones, MPa 8.0 / 5.7 / 4.0 8.0 / 5.5 / 4.0 8.0 / 5.5 / 4.0 8.0 / 5.5 / 4.0 Hydrogen-oil volume ratio at the reactor inlet 600:1 600:1 600:1 600:1 <![CDATA[Space velocity, h -1 > 1.3 1.3 1.3 1.3 Volume ratio of the first / second / third reaction zones, v% 60:20:20 50:30:20 55:25:20 50:25:25 Average reaction temperature in the first reaction zone, °C 342 347 345 345 Hydrogen sulfide concentration at the outlet of the first reaction zone, ppm 7500 7700 7500 7600 Average reaction temperature in the second reaction zone, °C 275 270 272 273 Average reaction temperature in the third reaction zone, °C 262 263 263 263 Liquid yield, % 98.5 98.2 98.4 98.35 Diesel sulfur content, ppm 7 7 8 8 Jet fuel mercaptan, ppm 9 9 9 8 Jet fuel sulfur content, ppm 880 855 865 860 Gasoline sulfur content, ppm 266 251 249 245 Gasoline olefin content, v% <1 <1 <1 <1

[0110] Table 3

[0111] Item Example 5 Example 6 Example 7 Example 8 First reaction zone FHUDS-8 FHUDS-8 FHUDS-8 FHUDS-8 Second reaction zone FH-40C FH-40C FH-40C FH-40C Third reaction zone FH-40B FH-40B FH-40B FH-40B Hydrogen partial pressure in the first / second / third reaction zones, MPa 7.0 / 4.5 / 3.5 7.0 / 4.0 / 3.5 8.0 / 5.7 / 4.0 8.0 / 5.7 / 4.0 Hydrogen-oil volume ratio at the reactor inlet 600:1 600:1 600:1 600:1 <![CDATA[Space velocity, h -1 > 1.3 1.3 1.3 1.3 Volume ratio of the first / second / third reaction zones, v% 60:20:20 60:20:20 60:20:20 60:20:20 Average reaction temperature in the first reaction zone, °C 342 342 342 342 Hydrogen sulfide concentration at the outlet of the first reaction zone, ppm 7500 7500 7500 7500 Average reaction temperature in the second reaction zone, °C 275 275 310 290 Average reaction temperature in the third reaction zone, °C 262 262 300 280 Liquid yield, % 98.63 98.65 97.85 98.02 Diesel sulfur content, ppm 9 9 7 8 Jet fuel mercaptan, ppm 9 9 2 5 Jet fuel sulfur content, ppm 955 996 205 596 Gasoline sulfur content, ppm 285 295 25 167 Gasoline olefin content, v% <1 <1 <1 <1

[0112] Table 4

[0113] Item Comparative Example 1 Comparative Example 2 Comparative Example 3 First reaction zone FHUDS-8 FHUDS-8 FHUDS-8 Second reaction zone FHUDS-8 FH-40C FH-40C Third reaction zone FHUDS-8 FH-40B FH-40B Hydrogen partial pressure in the first / second / third reaction zones, MPa 8.0 / 7.5 / 7.0 8.0 / 5.7 / 4.0 8.0 / 5.7 / 4.0 Hydrogen-oil volume ratio at the reactor inlet 600:1 600:1 600:1 <![CDATA[Space velocity, h -1 > 1.2 1.3 1.3 Volume ratio of the first / second / third reaction zones, v% 20:30:50 60:20:20 60:20:20 Average reaction temperature in the first reaction zone, °C 312 342 342 Hydrogen sulfide concentration at the outlet of the first reaction zone, ppm 3500 7500 7500 Average reaction temperature in the second reaction zone, °C 335 342 250 Average reaction temperature in the third reaction zone, °C 352 342 240 Liquid yield, % 97.3 97.5 98.6 Diesel sulfur content, ppm 6 6 7 Jet fuel mercaptan, ppm <1 <1 17 Jet fuel sulfur content, ppm <5 <5 2155 Gasoline sulfur content, ppm <1 <1 1680 Gasoline olefin content, v% <1 <1 2.2

[0114] It can be seen from the comparison between Example 1 and Examples 7 - 8 that by making the average reaction temperatures of the first, second, and third hydrogenation reactions be in the ranges of 340 - 400 °C, 260 - 300 °C, and 240 - 290 °C respectively, preferably in the ranges of 340 - 400 °C, 260 - 280 °C, and 240 - 270 °C, the sulfur contents of the obtained jet fuel and gasoline can be better controlled.

[0115] It can be seen from the data results of the above - mentioned examples and comparative examples that the most prominent feature of the hydrogenation method of the present invention is that by simply modifying a conventional coking gasoline and diesel hydrogenation refining unit and optimizing the catalyst grading method, through fractionation cutting, the reaction sequence and reaction process conditions of the reactants are deeply coupled, the reaction depth is controlled, the reaction heat is balanced, and the performance of the catalysts in each reaction zone is maximally exerted, so as to achieve the purpose of reducing the hydrogen consumption and energy consumption of the unit. The hydrogenation method of the present invention can increase the liquid yield and extend the operation cycle of the coking gasoline hydrogenation unit, bringing considerable economic and social benefits to enterprises and having great practical application advantages.

[0116] The preferred embodiments of the present invention have been described in detail above. However, the present invention is not limited thereto. Within the scope of the technical concept of the present invention, various simple modifications can be made to the technical solutions of the present invention, including any other suitable combination of each technical feature. These simple modifications and combinations should also be regarded as the content disclosed by the present invention and fall within the protection scope of the present invention.

Claims

1. A method for hydrogenating coking gasoline and diesel, characterized in that, The method comprises the following steps: (1) subjecting the coking gasoline and diesel raw oil to a first fractionation to obtain a light fraction of the raw oil, an intermediate fraction of the raw oil, and a heavy fraction of the raw oil; (2) in the presence of a first hydrogenation catalyst, subjecting the heavy fraction of the raw oil obtained in step (1) to a first hydrogenation reaction with hydrogen to obtain a first hydrogenation reaction product; (3) in the presence of a second hydrogenation catalyst, subjecting the first hydrogenation reaction product obtained in step (2) and the intermediate fraction of the raw oil obtained in step (1) to a second hydrogenation reaction to obtain a second hydrogenation reaction product; (4) in the presence of a third hydrogenation catalyst, subjecting the second hydrogenation reaction product obtained in step (3) and the light fraction of the raw oil obtained in step (1) to a third hydrogenation reaction to obtain a third hydrogenation reaction product; (5) subjecting the third hydrogenation reaction product obtained in step (4) to gas-liquid separation to obtain a gas phase and a liquid phase, and subjecting the liquid phase to stripping and a second fractionation to obtain a hydrogenated product; wherein, in step (2), the average reaction temperature of the first hydrogenation reaction is 340 - 420 °C; in step (3), the average reaction temperature of the second hydrogenation reaction is 260 - 310 °C; in step (4), the average reaction temperature of the third hydrogenation reaction is 240 - 300 °C, the average reaction temperature of the second hydrogenation reaction is 40 - 150 °C lower than that of the first hydrogenation reaction, and the average reaction temperature of the third hydrogenation reaction is 0 - 70 °C lower than that of the second hydrogenation reaction, the hydrogen partial pressure of the second hydrogenation reaction is 2 - 6 MPa lower than that of the first hydrogenation reaction, and the hydrogen partial pressure of the third hydrogenation reaction is 0 - 4 MPa lower than that of the second hydrogenation reaction, the first hydrogenation reaction in step (2) is mainly for hydrodesulfurization, the second hydrogenation reaction in step (3) is mainly for hydrodearomatization reaction and mild desulfurization reaction, and the third hydrogenation reaction in step (4) is mainly for hydrodesulfurization of mercaptan and desulfurization reaction, the hydrogen sulfide concentration in the product of the first hydrogenation reaction is 4000 - 14000 ppm.

2. The method according to claim 1, wherein In step (2), the conditions for the first hydrogenation reaction include: a hydrogen partial pressure of 3-16 MPa, a space velocity of 0.3-10 h -1 , and a hydrogen-oil volume ratio of 200:1 to 2500:

1.

3. The method according to claim 1, wherein, In step (2), the first hydrogenation catalyst comprises a first carrier and a first hydrogenation active component supported on the first carrier, and the first hydrogenation active component is W-Ni, Mo-Ni or W-Mo-Ni.

4. The method according to claim 3, wherein, The first carrier is one or more of alumina, silica, amorphous silica-alumina, and titanium oxide.

5. The method according to claim 3, wherein, Based on the weight of the first hydrogenation catalyst, the content of the first hydrogenation active component in terms of oxide is 15 - 50 wt%.

6. The method according to claim 5, wherein Based on the weight of the first hydrogenation catalyst, the content of the first hydrogenation active component in terms of oxide is 18 - 45 wt%.

7. The method according to claim 3, wherein The specific surface area of the first hydrogenation catalyst is 120 - 600 m 2 / g, and the pore volume is 0.2 - 1.5 mL / g.

8. The method according to claim 7, wherein, The specific surface area of the first hydrogenation catalyst is 200 - 450 m 2 / g, and the pore volume is 0.35 - 0.7 mL / g.

9. The method according to claim 1, wherein, In step (3), the conditions for the second hydrogenation reaction include: hydrogen partial pressure of 2 - 10 MPa, space velocity of 0.6 - 10 h -1 , and a hydrogen-to-oil volume ratio of 100:1 to 2500:

1.

10. The method according to claim 1, wherein, The second hydrogenation catalyst comprises a second carrier and a second hydrogenation active component supported on the second carrier, and the second hydrogenation active component is a metal oxide of Group VIB or / and a metal oxide of Group VIII.

11. The method according to claim 10, wherein, Based on the weight of the second hydrogenation catalyst, the content of the metal oxide of Group VIB is 5 - 30 wt%, and the content of the metal oxide of Group VIII is 1 - 15 wt%.

12. The method according to claim 11, wherein, Based on the weight of the second hydrogenation catalyst, the content of the metal oxide of Group VIB is 5 - 15 wt%, and the content of the metal oxide of Group VIII is 1.5 - 6 wt%.

13. The method according to claim 10, wherein, The specific surface area of the second hydrogenation catalyst is 100 - 550 m 2 / g, and the pore volume is 0.3 - 1.5 mL / g.

14. The method according to claim 13, wherein, The specific surface area of the second hydrogenation catalyst is 150 - 420 m 2 / g, and the pore volume is 0.4 - 0.9 mL / g.

15. The method according to claim 1, wherein, In step (4), the conditions for the third hydrogenation reaction include: hydrogen partial pressure of 1 - 8 MPa, space velocity of 0.6 - 10 h -1 , and a hydrogen-to-oil volume ratio of 100:1 to 2500:

1.

16. The method according to claim 15, wherein, The third hydrogenation catalyst includes a third carrier and a third hydrogenation active component supported on the third carrier, and the third hydrogenation active component is a Group VIB metal oxide or / and a Group VIII metal oxide.

17. The method according to claim 16, wherein, The Group VIII metal of the third hydrogenation active component is Co.

18. The method according to claim 16, wherein, Based on the weight of the third hydrogenation catalyst, the content of the Group VIB metal oxide is 5-30 wt%, and the content of the Group VIII metal oxide is 1-15 wt%.

19. The method according to claim 18, wherein Based on the weight of the third hydrogenation catalyst, the content of the Group VIB metal oxide is 5-15 wt%, and the content of the Group VIII metal oxide is 1.5-6 wt%.

20. The method according to claim 16, wherein The specific surface area of the third hydrogenation catalyst is 80 - 500 m 2 / g, and the pore volume is 0.3 - 1.5 mL / g.

21. The method according to claim 20, wherein, The specific surface area of the third hydrogenation catalyst is 100-380 m 2 / g, and the pore volume is 0.4-0.9 mL / g.

22. The method according to any one of claims 1-21, wherein, The average reaction temperature of the first hydrogenation reaction is 340-400 °C; the average reaction temperature of the second hydrogenation reaction is 260-310 °C; the average reaction temperature of the third hydrogenation reaction is 240-300 °C.

23. The method according to any one of claims 1-21, wherein The average reaction temperature of the second hydrogenation reaction is 50-100 °C lower than that of the first hydrogenation reaction, and the average reaction temperature of the third hydrogenation reaction is 5-50 °C lower than that of the second hydrogenation reaction.

24. The method according to any one of claims 1-21, wherein, The first hydrogenation reaction, the second hydrogenation reaction and the third hydrogenation reaction are carried out in the same reactor.

25. The method according to claim 24, wherein, The volume ratio of the first hydrogenation catalyst, the second hydrogenation catalyst and the third hydrogenation catalyst is (40-80):(10-40):(10-20).

26. The method according to claim 25, wherein, The volume ratio of the first hydrogenation catalyst, the second hydrogenation catalyst and the third hydrogenation catalyst is (75-50):(15-25):(10-25).

27. The method according to any one of claims 1-21, wherein, In step (1), the initial boiling point temperature of the coker gasoline and diesel feedstock is 30-100 °C, and the final boiling point temperature is 350-400 °C.

28. The method according to claim 27, wherein, The cut-off points of the light fraction, the middle fraction and the heavy fraction of the feedstock are 150-200 °C and 220-270 °C respectively.

29. The method according to any one of claims 1-21, wherein, This method further includes: after removing hydrogen sulfide from the gas phase obtained in step (5), returning it to step (1) to be mixed with the coker gasoline and diesel feedstock.

30. The method according to claim 29, wherein, The hydrogen sulfide concentration in the material of the gas phase obtained in step (5) after removing hydrogen sulfide is 0-1500 ppm.

31. The method according to claim 30, wherein, The hydrogen sulfide concentration in the material of the gas phase obtained in step (5) after removing hydrogen sulfide is 50-1000 ppm.

32. The method according to any one of claims 1-21, wherein, The hydrogenation products are gasoline, diesel and jet fuel.

33. The method according to claim 32, wherein, The sulfur content in the gasoline is 25-1500 ppm, and the sulfur content in the jet fuel is 500-1900 ppm.

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