A method of slurry bed hydrocracking

Abstract

This invention relates to the field of petrochemical energy technology and discloses a slurry-bed hydrocracking method. In the presence of a fluidized bed catalyst, biomass and / or plastics are hydrogenated with heavy oil in a fluidized bed reactor. The fluidized bed catalyst includes a support and active components. The active components include a first active component, a second active component, and a third active component. The first active component is selected from at least one element of Group IIIB, the second active component is selected from at least one element of Group IVB, and the third active component is selected from at least one element of Group VB. The products of the hydrocracking reaction are separated to obtain gas, gasoline and diesel fractions, and atmospheric residue. In the presence of the slurry-bed catalyst, the atmospheric residue is hydrocracking in the slurry-bed reactor. The fluidized bed catalyst of this invention, combined with the fluidized bed reactor, improves the overall conversion rate of the feedstock, the yield of gasoline and diesel, and reduces solid residue in the products, thereby extending the operating cycle of the equipment.

Description

Technical Field

[0001] This invention relates to the field of petrochemical energy technology, specifically to a method for slurry-bed hydrocracking. Background Technology

[0002] The proportion of low-quality heavy oil in refining feedstocks is increasing year by year, highlighting the urgent need for economical and efficient processing of this resource. Meanwhile, biomass energy, as an alternative energy source, is a crucial requirement for establishing a sustainable energy structure. However, current traditional biomass pyrolysis and catalytic pyrolysis technologies still face challenges such as low-quality conversion products and severe coking, limiting their large-scale utilization. Currently, waste plastic recycling in the industry typically employs physical and chemical recycling methods. However, traditional physical recycling technologies have significant technical shortcomings, failing to completely remove impurities and aged groups, thus hindering high-quality, high-value-added recycling. Chemical recycling, on the other hand, has a small industrial scale, facing problems such as high reaction energy consumption and difficulties in product separation and purification, making it difficult to achieve efficient recycling of waste plastics.

[0003] CN110028985A discloses a method for preparing high-quality fuel oil and / or chemical feedstock from biomass pyrolysis liquid. The method involves first subjecting the biomass pyrolysis liquid to hydrodeoxygenation in a fully mixed-flow catalyst circulation system within a fluidized bed reactor to obtain deoxygenated oil. However, due to the limited hydrorefining capacity of the fluidized bed reactor, the resulting deoxygenated oil still exhibits high oxygen content and coking reactivity. Therefore, the deoxygenated oil needs to be blended with at least one of heavy diesel oil, wax oil, and coal tar before it can be fed into a fixed-bed reactor for hydrocracking. It is evident that current hydrorefining processes adapted to heavy oil or heavy oil-biomass and / or plastic mixtures still face technical bottlenecks such as severe system coking and the need for further improvement in conversion rate and product quality.

[0004] Therefore, in order to make effective use of biomass and / or plastic resources and improve feedstock conversion rate and gasoline and diesel fraction yield, it is urgent to find a process that can simultaneously hydrogenate biomass and / or plastics as well as heavy oil. Summary of the Invention

[0005] The purpose of this invention is to overcome the above-mentioned problems existing in the prior art and to provide a slurry bed hydrocracking method.

[0006] To achieve the above objectives, the present invention provides a method for slurry-bed hydrocracking, the method comprising the following steps: (1) In the presence of a fluidized bed catalyst, biomass and / or plastics are subjected to a hydrogenation reaction with heavy oil in a fluidized bed reactor; wherein the fluidized bed catalyst comprises a support and an active component, the active component comprising a first active component, a second active component and a third active component, the first active component being selected from at least one element of Group IIIB, the second active component being selected from at least one element of Group IVB, and the third active component being selected from at least one element of Group VB; (2) Separate the products of the hydrogenation reaction to obtain gas, gasoline and diesel fractions and atmospheric residue; (3) In the presence of a slurry bed catalyst, the atmospheric residue oil obtained in step (2) is subjected to hydrocracking reaction in a slurry bed reactor.

[0007] The beneficial effects achieved by the present invention through the above technical solution are as follows: (1) In the presence of a specific fluidized bed catalyst, the present invention combines a fluidized bed reactor to hydrogenate a mixture of biomass and / or plastics and heavy oil, which can improve the total conversion rate of the raw materials, and the yield of gasoline and diesel is high, with less solid residue in the product, which can effectively reduce the risk of coking and blockage and improve the operating cycle of the equipment.

[0008] (2) In the prior art, the amount of biomass added is usually less than 20% of the total weight of the raw materials. Further increasing the amount of biomass added will cause severe coking of the equipment and prevent it from operating normally. However, the method of the present invention can increase the amount of biomass added to 50% of the total weight of the raw materials while ensuring the normal operation of the equipment, thereby achieving efficient utilization of biomass resources and further increasing the proportion of biomass replacing heavy oil.

[0009] (3) In a preferred embodiment, in addition to obtaining gasoline and diesel products with high yield, the method of the present invention can also obtain vacuum distillate oil, which is of high quality and can be directly used as a Group III or Group III+ lubricating base oil, and the added value of the product of the present invention is higher. Attached Figure Description

[0010] Figure 1 This is a flow chart of the combined process of biomass / waste plastics and heavy oil system slurry bed hydrocracking according to an embodiment of the present invention.

[0011] Explanation of reference numerals in the attached figures 1-Preheater for mixed feedstock, 2-Fluidized bed reactor, 3-Catalyst regeneration system, 4-Separation tower, 5-Slurry bed reactor, 6-Hot high-pressure separator, 7-Hot low-pressure separator, 8-Cold high-pressure separator, 9-Cold low-pressure separator, 10-Gas produced by fluidized bed reactor, 11-Gas and diesel fractions produced by fluidized bed reactor, 12-Ambient residue oil produced by fluidized bed reactor, 13-Slurry bed catalyst, 14-Gas produced by slurry bed reactor, 15-Gas and diesel fractions produced by slurry bed reactor, 16-Vacuum distillate oil, 17-Vacuum residue oil, 18-Reaction solids, 19-Regenerated catalyst. Detailed Implementation

[0012] The endpoints and any values ​​of the ranges disclosed herein are not limited to the precise ranges or values, and these ranges or values ​​should be understood to include values ​​close to these ranges or values. For numerical ranges, the endpoint values ​​of the various ranges, the endpoint values ​​of the various ranges and individual point values, and individual point values ​​can be combined with each other to obtain one or more new numerical ranges, which should be considered as specifically disclosed herein.

[0013] This invention provides a method for slurry-bed hydrocracking, the method comprising the following steps: (1) In the presence of a fluidized bed catalyst, biomass and / or plastics are subjected to a hydrogenation reaction with heavy oil in a fluidized bed reactor; wherein the fluidized bed catalyst comprises a support and an active component, the active component comprising a first active component, a second active component and a third active component, the first active component being selected from at least one element of Group IIIB, the second active component being selected from at least one element of Group IVB, and the third active component being selected from at least one element of Group VB; (2) Separate the products of the hydrogenation reaction to obtain gas, gasoline and diesel fractions and atmospheric residue; (3) In the presence of a slurry bed catalyst, the atmospheric residue oil obtained in step (2) is subjected to hydrocracking reaction in a slurry bed reactor.

[0014] This invention utilizes a fluidized bed catalyst in the presence of a specific fluidized bed catalyst and a fluidized bed reactor to hydrogenate a mixture of biomass and / or plastics with heavy oil. This process can improve the overall conversion rate of the feedstock, resulting in higher yields of gasoline and diesel, less solid residue in the products, effectively reducing the risk of coking and clogging, and extending the operating cycle of the equipment.

[0015] In this invention, the raw materials (biomass and / or plastics and heavy oil) can be directly fed into the fluidized bed reactor, or the raw materials can be mixed first (to form a homogeneous reaction slurry). Preferably, the raw materials are mixed before being fed into the fluidized bed reactor.

[0016] According to the present invention, preferably, the mixed raw materials are preheated in a mixed raw material preheater 1 before being fed into the fluidized bed reactor. More preferably, the preheating temperature is 300-400°C.

[0017] According to the present invention, preferably, based on the total weight of the biomass and / or plastics and heavy oil, the amount of biomass and / or plastics is 10-50% by weight (which can be any two values ​​from 10, 13, 16, 19, 22, 25, 28, 31, 34, 37, 40, 43, 45, 47, 50% by weight, or values ​​within that range), and the amount of heavy oil is 50-90% by weight (which can be any two values ​​from 50, 55, 60, 65, 70, 75, 80, 85, 90% by weight, or values ​​within that range).

[0018] In this invention, the composition of biomass and / or plastics and heavy oil is within the above-mentioned range, which can not only effectively avoid the reduction of heavy oil conversion rate, but also effectively increase the amount of biomass and / or plastics added, thereby reducing the demand for and dependence on petroleum resources.

[0019] More preferably, the method of the present invention can process raw materials with a biomass content of 20 wt% or more, and when the biomass content is 20 wt% or more, it can still effectively suppress coking in the reactor and pipeline blockage, maintain stable flow and continuous operation of the hydrogenation reaction system, and at the same time increase the conversion ratio of biomass to liquid fuel components, reduce the amount of heavy oil used, and achieve synergistic conversion of biomass and / or plastics and heavy oil.

[0020] According to the present invention, preferably, the heavy oil is at least one of atmospheric residue, vacuum residue, coal tar pitch and deasphalted oil, and more preferably coal tar pitch and / or vacuum residue.

[0021] According to the present invention, preferably, the initial boiling point of the heavy oil is not lower than 350℃ (preferably 350-550℃). In the present invention, the initial boiling point of the heavy oil is determined by the method of GB / T 9168-2025 "Determination of Distillation Range of Petroleum Products - Vacuum Distillation Method".

[0022] According to the present invention, preferably, the Concordant carbon content in the heavy oil is 8-55% by weight. The Concordant carbon content in the heavy oil is determined by the method of GB / T 268-2016 "Determination of Carbon Residue in Petroleum Products (Concordant Method)".

[0023] According to the present invention, preferably, the asphaltene content in the heavy oil is 5-30% by weight. The asphaltene is determined by the method of NB / SH / T 0509-2010 "Determination of Four Components of Petroleum Asphalt (Column Chromatography)".

[0024] According to the present invention, preferably, the sulfur content in the heavy oil is 0.5-6% by weight. The sulfur content is determined by the method of GB / T387-2012 "Determination of Sulfur Content in Dark Petroleum Products (Tube Furnace Method)".

[0025] According to the present invention, preferably, the nitrogen content in the heavy oil is 1000-10000 μg / g.

[0026] According to the present invention, preferably, the total metal (nickel, vanadium, iron) content in the heavy oil is 50-500 μg / g.

[0027] According to the present invention, preferably, the biomass is at least one of bamboo powder, wood chips, sawdust, bark, waste wood powder, pulp fiber, corn stalks, cotton stalks and rice husks, and more preferably at least one of bamboo powder, wood chips and sawdust.

[0028] According to the present invention, preferably, the biomass has a particle size of 60-80 mesh. In this invention, a particle size of 60-80 mesh refers to a composition that can pass through a 60-mesh sieve but not an 80-mesh sieve.

[0029] According to the present invention, preferably, the moisture content of the biomass is 3-12% by weight. The ash content of the biomass is 0.2-8% by weight. The volatile matter content of the biomass is 65-85% by weight. The fixed carbon content of the biomass is 10-25% by weight. The moisture, ash, volatile matter, and fixed carbon of the biomass are obtained by methods in GB / T 28731-2012 "Analytical Methods for Solid Biomass Fuels".

[0030] According to the present invention, preferably, the carbon content in the biomass is 42-52% by weight. The hydrogen content in the biomass is 5-7% by weight. The oxygen content in the biomass is 35-50% by weight. The nitrogen content in the biomass is 0.1-2% by weight. The sulfur content in the biomass is 0-0.5% by weight. In the present invention, the elemental content in the biomass is determined by the methods specified in GB / T 28734-2012 "Determination of Carbon and Hydrogen in Solid Biomass Fuels", GB / T 30728-2014 "Determination of Nitrogen in Solid Biomass Fuels", and GB / T 28732-2012 "Determination of Total Sulfur in Solid Biomass Fuels".

[0031] According to the present invention, preferably, the plastic is at least one selected from polyethylene, polypropylene, polystyrene, polyethylene terephthalate, polyvinyl chloride, acrylonitrile-butadiene-styrene copolymer and polyamide, and more preferably at least one selected from polyethylene, polypropylene and polystyrene.

[0032] According to the present invention, preferably, the weight-average molecular weight of the plastic is 1,000-500,000, more preferably 5,000-200,000. The particle size of the plastic is 40-100 mesh, more preferably 60-80 mesh. The softening temperature of the plastic is 80-260°C, more preferably 100-160°C.

[0033] According to the present invention, preferably, the plastic has a carbon content of 35-90% by weight, a hydrogen content of 3-16% by weight, an oxygen content of 0-35% by weight, a nitrogen content of 0-15% by weight, a sulfur content of 0-2% by weight, and a chlorine content of 0-60% by weight.

[0034] In this invention, the weight-average molecular weight of the plastic is determined by the method in GB / T 36214.1-2018 "Plastics - Determination of average molecular weight and molecular weight distribution of polymers by volume exclusion chromatography - Part 1: General rules"; the particle size of the plastic is determined by the method in GB / T 21843-2008 "Plastics - Determination of particle size by mechanical sieving of homopolymer and copolymer vinyl chloride resins"; and the softening temperature of the plastic is determined by the method in GB / T 1633-2025 "Plastics - Determination of Vicat softening temperature (VST) of thermoplastic plastics" or GB / T1634.2-2019 "Plastics - Determination of load deformation temperature - Part 2: Plastics and hard rubber".

[0035] According to the present invention, preferably, the fluidized bed reactor is a gas-liquid-solid three-phase fluidized bed reactor.

[0036] According to the present invention, preferably, the conditions for the hydrogenation reaction include: a reaction temperature of 370-470°C, a reaction pressure of 10-22 MPa, and a reaction volume hourly space velocity of 0.5-3 h⁻¹. -1 Hydrogen-to-oil volume ratio 500-2000 Nm 3 / m 3 .

[0037] In this invention, the hydrogenation reaction pressure refers to the pressure of the reaction system in the fluidized bed reactor, including the total pressure of hydrogen and the gases generated in the reaction; the hydrogenation reaction volume hourly space velocity refers to the total volume hourly space velocity of the total feedstock entering the fluidized bed reactor, including the volume hourly space velocity of biomass and / or plastics and heavy oil, as well as the volume hourly space velocity of the reflux vacuum residue.

[0038] According to the present invention, preferably, in step (2), the colloidal stability index of the atmospheric residue oil is 0.3-1.2, which can be any two values ​​formed by 0.3, 0.4, 0.5, 0.6, 0.7, 0.8, 0.9, 1, 1.1, 1.2, or values ​​within that range, and more preferably 0.3-0.8. When the colloidal stability index of the atmospheric residue oil is limited to the above range, the operating cycle of the device can be further improved.

[0039] In this invention, the colloidal stability index (CSI) of the atmospheric residue oil is obtained by the formula CSI = (saturated fraction content + asphaltenes content + quinoline insoluble content) / (aromatic fraction content + gum content + toluene insoluble content - quinoline soluble content). The contents of saturated fraction, aromatic fraction, gum, and asphaltenes are determined according to NB / SH / T 0509–2010 "Determination of Petroleum Asphalt Components"; the toluene insoluble content is determined according to GB / T 2292-2018 "Determination of Toluene Insoluble Content in Coking Products"; and the quinoline soluble content is determined according to GB / T 2293-2019 "Test Method for Quinoline Insoluble Content in Coking Asphalt Products". In this invention, the contents of all the above components are expressed as weight percent (%).

[0040] In this invention, the method for testing the viscosity of the atmospheric residue oil at 100°C is GB / T 265-1988 "Determination of Kinematic Viscosity and Calculation of Dynamic Viscosity of Petroleum Products". In this invention, the conditions for the hydrogenation reaction in step (1) and the separation in step (2) result in an atmospheric residue oil 12 with a viscosity of 300-750 mmHg at 100°C. 2 / s, further preferably 350-620mm 2 / s, more preferably 450-550mm 2 / s. When the viscosity of atmospheric residue oil 12 at 100°C is limited to the above range, the yield of gasoline and diesel fractions can be further improved, and the resulting vacuum distillate oil 16 can be directly used as a feedstock for Group III or Group III+ lubricating base oils.

[0041] According to the present invention, preferably, the first active component includes at least one of La, Ac, Ce, Sc, Y and Tb.

[0042] According to the present invention, preferably, the second active component includes at least one of Zr, Ti, Hf, Rf, Zr and Ti.

[0043] According to the present invention, preferably, the third active component includes at least one of V, Nb, Ta, Db, Nb and Ta.

[0044] According to the present invention, preferably, the active component is selected from at least one combination of La-Zr-V, Ac-Ti-Nb, Ce-Hf-Ta, Sc-Rf-Db, Y-Zr-Nb, and Tb-Ti-Ta. When the active component is selected from the above combinations, the overall conversion rate of the feedstock can be further improved, the yield of gasoline and diesel can be increased, and the yield of vacuum distillate oil can be increased. This vacuum distillate oil can be directly used as a feedstock for Group III or Group III+ lubricating base oils, thereby increasing the added value of the product.

[0045] According to the present invention, preferably, the support for the fluidized bed catalyst is alumina and / or silicon dioxide.

[0046] According to the present invention, preferably, based on the total weight of the fluidized bed catalyst, the content of the first active component is 0.1-8% by weight, more preferably 3-6% by weight; the content of the second active component is 0.5-15% by weight, more preferably 5-10% by weight; and the content of the third active component is 0.5-12% by weight, more preferably 6-10% by weight. In the present invention, the content of the active components in the fluidized bed catalyst within the above ranges is beneficial for further improving the raw material conversion rate and extending the operating cycle of the equipment.

[0047] According to the present invention, preferably, the specific surface area of ​​the fluidized bed catalyst is 250-400 m². 2 / g; the total pore volume of the fluidized bed catalyst is 0.6-1.5 cm³. 3 / g, the average pore size of the fluidized bed catalyst is 10-25 nm. In this invention, the specific surface area, total pore volume and average pore size of the catalyst are measured according to conventional nitrogen adsorption-desorption methods in the art.

[0048] The fluidized bed catalyst of the present invention can be purchased directly or prepared according to the methods described in the literature, for example, by impregnation. For example, the fluidized bed catalyst of the present invention can be prepared by the following method: 1) contacting a support (alumina and / or silica) with an aqueous solution of nitrates of the first, second, and third active components using an equal-volume impregnation method; 2) filtering, washing, and drying the mixture obtained in step 1); 3) mixing the solid obtained in step 2) with a binder (at least one of boehmite, alumina sol, silica sol, and aluminosilicate sol), extruding, drying, and calcining; 4) reducing under hydrogen. A fluidized bed catalyst with a particle size of 0.3-3 mm is obtained. The contact temperature in step 1) is 30-90℃ and the time is 1-24h; step 2) is dried at 80-150℃ for 2-24h; step 3) is dried at 80-150℃ for 2-24h; calcined at 350-650℃ for 1-10h; and step 4) is reduced at 250-550℃ for 1-10h.

[0049] The fluidized bed catalyst of the present invention can be impregnated using a shaped carrier, or the unshaped carrier can be impregnated first and then shaped using a binder.

[0050] In this invention, in step (2), the separation is carried out in separation tower 4. Preferably, the separation conditions include: a tower top pressure of 0.05-1 MPa, a tower top temperature of 40-180℃, a tower bottom temperature of 300-420℃, a reflux ratio of 0.1-5:1, and a theoretical number of trays of 5-50. In this invention, the gas 10 produced by the fluidized bed reactor is discharged from the top of separation tower 4, the gasoline and diesel fraction 11 is collected from the top and / or side stream of separation tower 4, and the atmospheric residue oil 12 is collected from the bottom of separation tower 4.

[0051] In this invention, the separation refers to separating the hydrogenation reaction products to obtain gas, gasoline and diesel fractions and atmospheric residue oil, wherein the gas (gas 10 produced by the fluidized bed reactor) includes components such as CO, H2S, and C1-C5 hydrocarbons.

[0052] According to the present invention, preferably, step (1) further includes discharging the unconverted material, deactivated fluidized bed catalyst, and coking material from the fluidized bed reactor, wherein the deactivated fluidized bed catalyst is returned to the fluidized bed reactor after being treated by the catalyst regeneration system 3. In the present invention, a portion of the fluidized bed catalyst and the unconverted material are discharged from the fluidized bed reactor together, and the deactivated fluidized bed catalyst is regenerated in the catalyst regeneration system 3. The method for regenerating the deactivated fluidized bed catalyst includes calcining the deactivated fluidized bed catalyst in air and then reducing it under hydrogen, wherein the calcination conditions include a temperature of 350-650°C and a time of 1-10 h; the reduction conditions include a temperature of 250-550°C and a time of 1-10 h.

[0053] According to the present invention, preferably, the slurry bed reactor is a fully backmixed reactor or an empty tank reactor capable of internal and external circulation.

[0054] According to the present invention, preferably, the slurry bed catalyst is an oil-soluble dispersion catalyst.

[0055] According to the present invention, preferably, the slurry bed catalyst comprises an oil-soluble organic ligand, a first metal component, and a second metal component; more preferably, the first metal component comprises at least one of Sn, Tb, and Ta; and the second metal component comprises one or more of Fe, W, Sn, Tb, Ta, Mo, and Ni.

[0056] According to the present invention, preferably, the oil-soluble organic ligand comprises cycloalkanoic acid, isooctanoic acid, 2-ethylhexanoic acid, neodecanoic acid, oleic acid, stearic acid, acetylacetone, dialkyl dithiocarbamic acid, dialkyl dithiophosphate, xanthic acid, and C6-C... 30 At least one of the organic amines, more preferably at least one of cycloalkanoic acid, isooctanoic acid, 2-ethylhexanoic acid, dialkyl dithiocarbamic acid and dialkyl dithiophosphoric acid.

[0057] According to the present invention, preferably, based on the total weight of the slurry bed catalyst, the content of the first metal component is 0.01-15% by weight; and the content of the second metal component is 1-40% by weight.

[0058] According to the present invention, preferably, the second metal component comprises Mo and Ni; more preferably, the weight ratio of Mo to Ni is 1:0.8-1.2. In the present invention, the types and contents of the components in the slurry bed catalyst 13 are within the above-mentioned range, which is beneficial for further improving the hydrocracking effect.

[0059] The slurry bed catalyst 13 of the present invention can be purchased directly or prepared according to the methods described in the literature. For example, the slurry bed catalyst 13 of the present invention can be prepared by the following method: mixing an oil-soluble organic ligand with an organic solvent (e.g., xylene) and stirring at 20-150°C for 0.5-8 h to obtain a ligand solution; then adding a first metal component precursor and a second metal component precursor to the ligand solution and performing a complexation reaction at 40-180°C for 0.5-12 h to obtain a reaction solution containing a metal-organic complex; allowing the reaction solution to stand for 1-5 h to separate into layers, removing the aqueous phase, and washing, dehydrating (dehydrating at 100-130°C for 1-5 h) and filtering the organic phase to obtain an oil-soluble metal-organic complex; optionally, pre-sulfurizing the oil-soluble metal-organic complex with a sulfurizing agent at 100-350°C for 0.5-8 h to obtain the slurry bed catalyst.

[0060] According to the present invention, preferably, the vulcanizing agent used for pre-vulcanization is at least one selected from dimethyl disulfide, carbon disulfide, sulfur powder, thiocycloalkanoic acid, diethyl disulfide, dibenzyl disulfide, n-butanethiol and polysulfides.

[0061] According to the present invention, preferably, the molar ratio of sulfur to the total amount of Sn, Mo and Ni metal elements in the dimethyl disulfide is 3-6:1.

[0062] According to the present invention, preferably, the hydrocracking reaction conditions include: a reaction temperature of 380-520°C, a reaction pressure of 15-30 MPa, and a reaction volume hourly space velocity of 0.1-2 h⁻¹. -1 Hydrogen-to-oil volume ratio 200-1000 Nm 3 / m3 .

[0063] In this invention, the hydrocracking reaction pressure refers to the total reaction pressure in the slurry bed reactor; the hydrocracking reaction volume hourly space velocity refers to the volume hourly space velocity of the atmospheric residue oil 12 obtained from the bottom of the separation tower 4 entering the slurry bed reactor.

[0064] According to the present invention, preferably, the method further includes step (4), separating the product of step (3) to obtain gas, gasoline and diesel fractions, vacuum distillate oil and vacuum residue oil; wherein the vacuum residue oil is returned to the fluidized bed reactor for recycling reaction.

[0065] According to the present invention, preferably, in step (4), the separation is carried out by a hot high-pressure separator, a hot low-pressure separator, a cold high-pressure separator and a cold low-pressure separator.

[0066] According to the present invention, preferably, the separation temperature of the hot high-pressure separator is 300-450℃, the separation pressure is 8-25MPa, and the material residence time is 1-60min.

[0067] According to the present invention, preferably, the separation temperature of the hot low-pressure separator is 250-430℃, the separation pressure is 0.01-3MPa, and the material residence time is 1-90min. In the present invention, vacuum distillate oil is obtained at the top and / or side of the hot low-pressure separator, and vacuum residue oil is obtained at the bottom.

[0068] According to the present invention, preferably, the separation temperature of the cold high-pressure separator is 20-120℃, the separation pressure is 8-25MPa, and the material residence time is 1-60min. In the present invention, gas is obtained at the top of the cold high-pressure separator and liquid phase product is obtained at the bottom.

[0069] According to the present invention, preferably, the separation temperature of the cold low-pressure separator is 20-120℃, the separation pressure is 0.01-3MPa, and the material residence time is 1-90min, and gasoline and diesel fractions 15 are obtained at the bottom of the cold low-pressure separator 9.

[0070] In this invention, the separating reactors are a hot high-pressure separator, a hot low-pressure separator, a cold high-pressure separator, and a cold low-pressure separator. The products of the hydrocracking reaction in the slurry bed reactor 5 are separated by the hot high-pressure separator 6. The product obtained at the top of the hot high-pressure separator 6 enters the cold high-pressure separator 8, and the product obtained at the bottom of the hot high-pressure separator 6 enters the hot low-pressure separator 7. The liquid phase product obtained at the bottom of the cold high-pressure separator 8 enters the cold low-pressure separator 9, and the product obtained at the top of the cold high-pressure separator 8 is the gas 14 produced by the slurry bed reactor. The vacuum residue oil 17 is obtained at the bottom of the hot low-pressure separator 7, and the vacuum distillate oil 16 is obtained at the top and / or side of the hot low-pressure separator 7. The gasoline and diesel fractions 15 produced by the slurry bed reactor are obtained by the cold low-pressure separator 9. The gas 14 produced by the slurry bed reactor includes components such as CO, H2S, and C1-C5 hydrocarbons.

[0071] According to the present invention, preferably, the ratio of the vacuum residue oil returned to the fluidized bed reactor to the fresh feedstock of the fluidized bed in step (1) is 10-20 by weight.

[0072] According to the present invention, preferably, the gasoline and diesel fractions (including gasoline and diesel fraction 11 produced by the fluidized bed reactor and gasoline and diesel fraction 15 produced by the slurry bed reactor) are fed into a refining reactor for refining.

[0073] according to Figure 1 The process flow diagram shown illustrates the combined slurry-bed hydrocracking process for a biomass and / or plastics and heavy oil system. The specific steps for the hydrocracking reaction are as follows: (1) The biomass is crushed to 60-80 mesh (and / or the plastic is crushed to 60-80 mesh) and mixed evenly with heavy oil. The above slurry is then heated by the mixing raw material preheater 1 (temperature is 300-400℃) and injected from the bottom of the fluidized bed reactor 2 for hydrogenation reaction. Unconverted biomass residue (and / or plastic residue) and reaction coking material and other reaction solids 18 are discharged from the bottom side stream of the fluidized bed reactor 2. The hydrogenation reaction product enters the separation tower 4 from the top of the fluidized bed reactor 2.

[0074] (2) The gasoline and diesel fraction 11 produced by the fluidized bed reactor is collected from the top and / or side stream of the separation tower 4, and the gas 10 produced by the fluidized bed reactor is collected from the top of the separation tower 4. The atmospheric residue oil 12 (boiling point > 350℃) produced by the fluidized bed reactor obtained from the bottom of the separation tower 4 is sent to the slurry bed reactor 5. In addition, the deactivated catalyst discharged from the bottom side stream of the fluidized bed reactor 2 is regenerated by the catalyst regeneration system 3 to obtain the regenerated catalyst 19, which is then recycled back to the fluidized bed reactor 2.

[0075] (3) The atmospheric residue oil 12 (boiling point > 350℃) produced by the fluidized bed reactor is mixed with the slurry bed catalyst 13 and then enters the reactor from the bottom of the slurry bed reactor 5 for hydrocracking reaction.

[0076] (4) The products of the hydrocracking reaction in the slurry bed reactor are fed into the hot high-pressure separator 6 for separation. The product obtained from the top of the hot high-pressure separator 6 enters the cold high-pressure separator 8, and the product obtained from the bottom of the hot high-pressure separator 6 enters the hot low-pressure separator 7. The liquid phase product obtained at the bottom of the cold high-pressure separator 8 enters the cold low-pressure separator 9, and the product obtained at the top of the cold high-pressure separator 8 is the gas 14 produced by the slurry bed reactor. The vacuum residue oil 17 is obtained at the bottom of the hot low-pressure separator 7, and the vacuum distillate oil 16 is obtained at the top and / or side of the hot low-pressure separator 7. The vacuum residue oil 17 can be mixed with fresh feedstock and then enters the fluidized bed reactor 2. The gasoline and diesel fraction 15 produced by the slurry bed reactor is obtained at the bottom of the cold low-pressure separator 9. The gasoline and diesel fraction 15 and the gasoline and diesel fraction 11 produced by the fluidized bed reactor separated by the separation tower 4 are sent together to the subsequent refining reactor to finally obtain qualified clean gasoline and diesel products.

[0077] The present invention will be described in detail below through embodiments.

[0078] The polypropylene waste plastic has a particle size of 70 mesh, a weight-average molecular weight of 52,000, a softening temperature of 136℃, and a density of 0.9 g / cm³. 3 The carbon content is 85.4 wt%, the hydrogen content is 14.1 wt%, the oxygen content is 0.35 wt%, the nitrogen content is 0.08 wt%, the sulfur content is 0.02 wt%, and the chlorine content is 0.05 wt%.

[0079] The composition and properties of bamboo powder and vacuum residue feedstock are shown in Tables 1 and 2: Table 1

[0080] Table 2

[0081] The fluidized bed catalysts used in the examples and comparative examples were prepared by loading them onto a support using a saturated impregnation method. The specific method is as follows: 1) The support (alumina) was contacted with an aqueous solution of nitrates from the first, second, and third active components using an equal-volume impregnation method; 2) The mixture obtained in step 1) was filtered, washed, and dried; 3) The solid obtained in step 2) was mixed with a binder (boehmite), extruded, dried, and calcined; 4) Reduction was carried out under hydrogen. A fluidized bed catalyst with a particle size of 2 mm was obtained. The contact temperature in step 1) was 70°C for 10 h; the drying in step 2) was at 110°C for 8 h; the drying in step 3) was at 120°C for 6 h; the calcination in step 3) was at 550°C for 5 h; and the reduction in step 4) was at 360°C for 4 h.

[0082] The specific preparation method of the slurry bed catalyst used in the examples and comparative examples is as follows: 100 parts by weight of naphthenic acid and 120 parts by weight of xylene were added to a reactor, and the mixture was heated to 70°C under nitrogen protection and stirred for 1.5 h to obtain a ligand solution; 5 parts by weight of stannous chloride, 40 parts by weight of ammonium heptamolybdate, and 25 parts by weight of nickel nitrate were dissolved in 100 parts by weight of deionized water to obtain an aqueous solution of the metal precursor; the aqueous solution of the metal precursor was added dropwise to the ligand solution, and after the addition was completed, the pH of the system was adjusted to 7.5 with ammonia water, and then the temperature was raised to 100°C and the complexation reaction was stirred for 6 h to obtain a Sn-containing solution. The reaction solution of the Sn-Mo-Ni metal-organic complex was prepared; the reaction solution was allowed to stand for 2 hours to separate into layers, the aqueous phase was removed, the obtained organic phase was washed three times with deionized water, then dehydrated at 120°C for 3 hours, and then filtered to remove insoluble matter to obtain an oil-soluble Sn-Mo-Ni metal-organic complex; the oil-soluble Sn-Mo-Ni metal-organic complex was mixed with dimethyl disulfide, wherein the molar ratio of sulfur in dimethyl disulfide to the total amount of Sn, Mo and Ni metal elements was 4:1, and the mixture was heated to 240°C under a nitrogen atmosphere and contacted for 3 hours for pre-sulfurization to obtain the slurry bed catalyst.

[0083] The formula for calculating the total conversion rate is: (m 步骤（2）中气体组分10 +m 步骤（2）汽柴油馏分11 +m 步骤（4）中气体14 +m 步骤（4）汽柴油馏分15 +m 减压馏分油 ) / m 原料总重 ×100% by weight% The formula for calculating the yield of vacuum distillate oil is: m 减压馏分油 / m 原料 ×100% by weight% The formula for calculating the yield of gasoline and diesel is: m 汽柴油馏分 / m 原料 ×100% by weight% The formula for calculating the oxygen removal rate of biomass is: [m 生物质 ×w 生物质-(m 汽柴油馏分 ×w 汽柴油馏分 +m 减压馏分油 ×w 减压馏分油 +m 固体残渣 ×w 固体残渣 )] / (m 生物质 ×w 生物质 W represents the oxygen content in each component.

[0084] Example 1 use Figure 1 The process flow shown is as follows: (1) Bamboo powder pulverized to 60-80 mesh is mixed with coal tar pitch at a weight ratio of 30:70. After being heated to 320°C in the preheater 1, the mixture is injected from the bottom into the fluidized bed reactor 2 for hydrogenation reaction. The fluidized bed catalyst is an Ac-Ti-Nb / Al2O3 catalyst (containing 4wt% Ac, 6wt% Ti, 7wt% Nb, and 83wt% Al2O3, with a specific surface area of ​​350 m²). 2 / g, total pore volume is 1.2cm³ 3 / g, average pore size is 20nm); the operating conditions of the fluidized bed reactor are: reaction temperature 420℃, reaction pressure 20MPa, hydrogen-to-oil volume ratio 500Nm 3 / m 3 Volumetric space velocity 1h -1 .

[0085] (2) The reaction products enter the separation tower 4 from the top of the fluidized bed reactor 2. A top component is obtained at the top of the separation tower 4. After cooling to room temperature, the top component yields gasoline and diesel fraction 11 and gas 10 produced by the fluidized bed reactor. The atmospheric residue oil 12 (boiling point above 350℃) obtained from the bottom of the separation tower 4 is sent to the slurry bed reactor 5. The deactivated catalyst discharged from the bottom side stream of the fluidized bed reactor 2 is regenerated by the catalyst regeneration system 3 to obtain regenerated catalyst 19, which is then recycled back to the fluidized bed reactor 2. The deactivated fluidized bed catalyst is calcined in air and then reduced under hydrogen. The calcination conditions include a temperature of 630℃ and a time of 8 hours; the reduction conditions include a temperature of 500℃ and a time of 6 hours. Additionally, solid materials 18, such as biomass residues and reaction coke, are discharged from the bottom side stream of the fluidized bed reactor 2.

[0086] (3) After mixing atmospheric residue 12 and oil-soluble dispersed Sn-Mo-Ni-based catalyst (containing 2wt% Sn, 8wt% Mo, 7wt% Ni, and 83wt% organic ligand (cycloalkanoic acid)), the mixture is introduced into the slurry bed reactor 5 from the bottom for hydrocracking. The operating conditions of the slurry bed reactor are: reaction temperature 440℃, reaction pressure 18MPa, and hydrogen-to-oil volume ratio 650Nm. 3 / m 3 Volumetric space velocity 1h -1 .

[0087] (4) The products of the hydrocracking reaction in the slurry bed reactor are fed into the hot high-pressure separator 6 (the conditions are separation temperature of 350℃, separation pressure of 20MPa, and residence time of 20min) for separation. The products obtained at the top of the hot high-pressure separator 6 enter the cold high-pressure separator 8, and the products obtained at the bottom of the hot high-pressure separator 6 enter the hot low-pressure separator 7. The products obtained at the bottom of the cold high-pressure separator 8 (the conditions are separation temperature of 100℃, separation pressure of 20MPa, and residence time of 60min) enter the cold low-pressure separator 9, and the products obtained at the top of the cold high-pressure separator 8 are gas 14. The vacuum residue 17 and vacuum distillate 16 obtained by the hot low-pressure separator 7 (the conditions are separation temperature of 350℃, separation pressure of 0.3MPa, and residence time of 20min) are mixed with fresh feed and then fed into the fluidized bed reactor 2. The reflux vacuum residue is about 18wt% of the fresh feed. The gasoline and diesel fraction 15 from the slurry bed reactor is obtained by separation in the cold low-pressure separator 9 (conditions: separation temperature 60℃, separation pressure 0.3MPa, residence time 30min). This gasoline and diesel fraction 15, along with the gasoline and diesel fraction 11 from the slurry bed reactor separated in the separation tower 4, are fed into the subsequent refining reactor to finally obtain qualified clean gasoline and diesel products. The parameters during the initial operation period (5 days) are shown in Table 3. The continuous operation time is defined as the duration of slurry bed operation when the total conversion rate decreases to 90% of the initial conversion rate.

[0088] Example 2 The method described in Example 1 was followed, except that bamboo powder and coal tar pitch were mixed at a weight ratio of 50:50; the operating conditions of the fluidized bed reactor were: reaction temperature 430℃, reaction pressure 20MPa, and hydrogen-to-oil volume ratio 1500 Nm³. 3 / m 3 Volume hourly space velocity 1.5 h⁻¹ -1 Slurry bed reactor operating conditions: reaction temperature 460℃, reaction pressure 22MPa, hydrogen-to-oil volume ratio 700Nm³. 3 / m 3 Volumetric space velocity 0.5 h⁻¹ -1The reflux vacuum residue is approximately 15 wt% of the fresh feedstock.

[0089] Example 3 The method described in Example 1 was followed, except that bamboo powder and coal tar pitch were mixed at a weight ratio of 15:85; the operating conditions of the fluidized bed reactor were: reaction temperature 440℃, reaction pressure 14MPa, and hydrogen-to-oil volume ratio 1800 Nm³. 3 / m 3 Volumetric space velocity 2h -1 Slurry bed reactor operating conditions: reaction temperature 470℃, reaction pressure 24MPa, hydrogen-to-oil volume ratio 1000Nm³. 3 / m 3 Volumetric space velocity 1h -1 The reflux vacuum residue is approximately 13 wt% of the fresh feedstock.

[0090] Example 4 The method described in Example 1 was followed, except that bamboo powder was replaced with waste polypropylene plastic.

[0091] Example 5 The method described in Example 1 was followed, except that the fluidized bed catalyst was Ac-Zr-V / Al2O3 (containing 10wt% Ac, 18wt% Zr, 14wt% V and 58wt% Al2O3, with a specific surface area of ​​280 m²). 2 / g, total pore volume is 0.8cm³ 3 / g, average pore size 15nm).

[0092] Example 6 The method described in Example 1 was followed, except that bamboo powder and coal tar pitch were mixed at a weight ratio of 55:45; the operating conditions of the fluidized bed reactor were: reaction temperature 420°C, reaction pressure 15 MPa, and hydrogen-to-oil volume ratio 600 Nm³. 3 / m 3 Volumetric space velocity 0.5 h⁻¹ -1 The reflux vacuum residue is approximately 9 wt% of the fresh feedstock.

[0093] Example 7 The method described in Example 1 is followed, except that the vacuum residue separated by the hot low-pressure separator 7 is discharged from the bottom of the reaction system and is not recycled back to the fluidized bed reactor.

[0094] Example 8 The method described in Example 1 is followed, except that the vacuum residue 11 obtained by the hot low-pressure separator 7 is not returned to the fluidized bed reactor 2, but is mixed with the atmospheric residue 12 and then enters the slurry bed reactor. The amount of vacuum residue returned is about 18 wt% of the atmospheric residue.

[0095] Example 9 The procedure was carried out according to the method described in Example 1, except that the volume hourly space velocity was maintained at 0.5 h. -1 Under conditions of higher temperature (450℃) and higher reaction pressure (23MPa) for fluidized bed hydrogenation and separation, the viscosity of atmospheric residue oil at 100℃ was reduced to 330 mmHg. 2 / s.

[0096] Comparative Example 1 The method described in Example 8 is followed, except that the fluidized bed reactor is not included. Instead, the mixed feedstock is heated to 320°C in the mixed feedstock preheater 1 and then injected from the bottom into the slurry bed reactor for hydrocracking (the vacuum residue discharged from the bottom of the hot low-pressure separator 7 is recycled back to the slurry bed reactor).

[0097] Comparative Example 2 The method described in Example 8 differs from that described in Example 8 in that the fluidized bed catalyst is Ni-Mo / Al2O3 (containing 15wt% Mo, 8wt% Ni and 77wt% Al2O3, with a specific surface area of ​​300 m²). 2 / g, total pore volume is 0.85cm³ 3 / g, average pore size 14nm).

[0098] Comparative Example 3 The method described in Example 8 was followed, except that the slurry-bed reactor was replaced with a fixed-bed reactor. The conditions of the fixed-bed reactor were a reaction temperature of 400°C, a reaction pressure of 18 MPa, and a volume hourly space velocity of 0.6 h⁻¹. -1 The hydrogen-to-oil ratio is 1000 Nm³ / m³.

[0099] Comparative Example 4 The method described in Example 8 is followed, except that the fluidized bed reactor is replaced with a slurry bed reactor, and the operating conditions of the slurry bed reactor are the same as those of the fluidized bed reactor in Example 8.

[0100] Table 3

[0101] Note: " / " indicates that it is the same as in Example 1. "-" indicates that the parameter is not present.

[0102] QI refers to quinoline insolubles, which are determined by the method specified in GB / T 2293-2019 "Determination of Quinoline Insolubles in Coking Pitch Products".

[0103] MCR refers to the micro carbon residue value, which is determined by the micro method in GB / T 17144-2021 "Determination of carbon residue in petroleum products - micro method".

[0104] Table 4

[0105] Note: " / " indicates that it is the same as in Example 1; "-" indicates that the parameter is not present.

[0106] QI refers to quinoline insolubles, which are determined by the method specified in GB / T 2293-2019 "Determination of Quinoline Insolubles in Coking Pitch Products".

[0107] MCR refers to the micro carbon residue value, which is determined by the micro method in GB / T 17144-2021 "Determination of carbon residue in petroleum products - micro method".

[0108] The data related to the fluidized bed discharge in Comparative Example 4 refers to the data related to the discharge of the slurry bed after the slurry bed was replaced.

[0109] Test case The vacuum distillate oils obtained in the examples and comparative examples were tested for saturated hydrocarbon content, sulfur content, and viscosity index VI. The test results are shown in Table 5. The saturated hydrocarbon content was determined using the method specified in NB / SH / T 0885-2014 "Determination of Total Aromatic Hydrocarbons and Total Saturated Hydrocarbons in Lubricating Oil Base Oils - Differential Refractive Index Detector High Performance Liquid Chromatography"; the sulfur content was determined using the method specified in GB / T17040-2019 "Determination of Sulfur Content in Petroleum and Petroleum Products - Energy Dispersive X-ray Fluorescence Spectrometry"; and the viscosity index VI was determined using the method specified in GB / T 1995-1998 "Calculation Method for Viscosity Index of Petroleum Products".

[0110] Table 5

[0111] As can be seen from the results in Tables 3-5, the raw materials (biomass and / or plastics and heavy oil) provided by this invention, combined with fluidized bed hydrotreating for raw material pretreatment and slag removal, will slow down the deactivation rate of the slurry bed hydrotreating catalyst and reduce coking, thereby improving the total conversion rate and the yield of gasoline and diesel fractions and increasing the operating cycle of the equipment.

[0112] Replacing the catalyst in the fluidized bed reaction with a common Mo-Ni-based catalyst (Comparative Example 2) results in the atmospheric residue oil from the fluidized bed being less suitable for subsequent deep conversion in the slurry bed, ultimately shortening the equipment's operating cycle. This is accompanied by a decrease in total conversion rate and gasoline / diesel fraction yield, and the vacuum distillate oil is also difficult to use as a lubricating oil base stock. The properties required for Group III or Group III+ lubricating base stocks are as follows: saturated fraction ≥ 90%, sulfur content ≤ 0.03 wt%, viscosity index VI ≥ 120.

[0113] Replacing the slurry-bed hydrogenation reactor in the combined process described in this invention with a fixed-bed hydrogenation reactor (Comparative Example 3) significantly shortens the continuous operation time, resulting in a limited final conversion depth.

[0114] Replacing the fluidized bed reactor in the combined process described in this invention with a slurry bed hydrogenation reactor (Comparative Example 4) resulted in severe catalyst deactivation, which deteriorated the properties of the circulating residue oil, ultimately leading to a reduction in conversion rate, severe equipment coking, and a significant reduction in equipment operating cycle.

[0115] 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 inventive concept, various simple modifications can be made to the technical solutions of the present invention, including combinations of various technical features in any other suitable manner. These simple modifications and combinations should also be considered as the content disclosed in the present invention and are all within the protection scope of the present invention.

Claims

1. A method for slurry-bed hydrocracking, characterized in that, The method includes the following steps: (1) In the presence of a fluidized bed catalyst, biomass and / or plastics are subjected to a hydrogenation reaction with heavy oil in a fluidized bed reactor; wherein the fluidized bed catalyst comprises a support and an active component, the active component comprising a first active component, a second active component and a third active component, the first active component being selected from at least one element of Group IIIB, the second active component being selected from at least one element of Group IVB, and the third active component being selected from at least one element of Group VB; (2) Separate the products of the hydrogenation reaction to obtain gas, gasoline and diesel fractions and atmospheric residue; (3) In the presence of a slurry bed catalyst, the atmospheric residue oil obtained in step (2) is subjected to hydrocracking reaction in a slurry bed reactor.

2. The method according to claim 1, wherein, Based on the total weight of the biomass and / or plastics and heavy oil, the amount of biomass and / or plastics used is 10-50% by weight, and the amount of heavy oil used is 50-90% by weight. And / or, the heavy oil is at least one of atmospheric residue, vacuum residue, coal tar pitch and deasphalted oil, preferably coal tar pitch and / or vacuum residue; And / or, the initial boiling point of the heavy oil is not lower than 350°C, and the content of KORP in the heavy oil is 8-55% by weight, the content of asphaltenes is 5-30% by weight, and the sulfur content is 0.5-6% by weight; And / or, the biomass is at least one of bamboo powder, wood chips, sawdust, bark, waste wood powder, pulp fiber, corn stalks, cotton stalks and rice husks; And / or, the biomass has a particle size of 60-80 mesh; the biomass contains 3-12% moisture, 0.2-8% ash, 65-85% volatile matter, and 10-25% fixed carbon; the biomass contains 42-52% carbon, 5-7% hydrogen, 35-50% oxygen, 0.1-2% nitrogen, and 0-0.5% sulfur. And / or, the plastic is at least one of polyethylene, polypropylene, polystyrene, polyethylene terephthalate, polyvinyl chloride, acrylonitrile-butadiene-styrene copolymer and polyamide; And / or, the weight-average molecular weight of the plastic is 1,000-500,000; the particle size of the plastic is 40-100 mesh; and the softening temperature is 80-260℃.

3. The method according to claim 1 or 2, wherein, The conditions for the hydrogenation reaction include: a reaction temperature of 370-470℃, a reaction pressure of 10-22 MPa, and a reaction volume hourly space velocity of 0.5-3 h⁻¹. -1 Hydrogen-to-oil volume ratio 500-2000 Nm 3 / m 3 .

4. The method according to any one of claims 1-3, wherein, In step (2), the colloidal stability index of the atmospheric residue oil is 0.3-1.2, preferably 0.3-0.8; And / or, the conditions of the hydrogenation reaction described in step (1) and the separation described in step (2) result in an atmospheric residue oil viscosity of 300-750 mm at 100°C. 2 / s, preferably 350-620mm 2 / s.

5. The method according to any one of claims 1-4, wherein, The first active component includes at least one of La, Ac, Ce, Sc, Y, and Tb; The second active component includes at least one of Zr, Ti, Hf, Rf, Zr, and Ti; The third active component includes at least one of V, Nb, Ta, Db, Nb, and Ta; And / or, the support for the fluidized bed catalyst is alumina and / or silica; Preferably, based on the total weight of the fluidized bed catalyst, the content of the first active component is 0.1-8% by weight; the content of the second active component is 0.5-15% by weight; and the content of the third active component is 0.5-12% by weight. More preferably, the active component is selected from at least one combination of La-Zr-V, Ac-Ti-Nb, Ce-Hf-Ta, Sc-Rf-Db, Y-Zr-Nb and Tb-Ti-Ta.

6. The method according to any one of claims 1-5, wherein, In step (2), the separation conditions include: the pressure at the top of the separation tower is 0.05-1 MPa, the temperature at the top of the tower is 40-180℃, the temperature at the bottom of the tower is 300-420℃, the reflux ratio is 0.1-5:1, and the theoretical number of plates is 5-50.

7. The method according to any one of claims 1-6, wherein, Step (1) also includes discharging the unconverted material, deactivated fluidized bed catalyst, and coking material from the fluidized bed reactor, wherein the deactivated fluidized bed catalyst is regenerated and returned to the fluidized bed reactor.

8. The method according to any one of claims 1-7, wherein, The slurry bed catalyst is an oil-soluble dispersion catalyst; Preferably, the slurry bed catalyst comprises an oil-soluble organic ligand, a first metal component, and a second metal component; more preferably, the first metal component comprises at least one of Sn, Tb, and Ta; and the second metal component comprises one or more of Fe, W, Sn, Tb, Ta, Mo, and Ni. More preferably, the oil-soluble organic ligand includes cycloalkanoic acid, isooctanoic acid, 2-ethylhexanoic acid, neodecanoic acid, oleic acid, stearic acid, acetylacetone, dialkyl dithiocarbamic acid, dialkyl dithiophosphate, xanthic acid, and C6-C... 30 At least one of the organic amines; More preferably, based on the total weight of the slurry bed catalyst, the content of the first metal component is 0.01-15% by weight; and the content of the second metal component is 1-40% by weight.

9. The method according to any one of claims 1-8, wherein, The hydrocracking reaction conditions include: reaction temperature 380-520℃, reaction pressure 15-30MPa, and reaction volume hourly space velocity 0.1-2h. -1 Hydrogen-to-oil volume ratio 200-1000 Nm 3 / m 3 .

10. The method according to any one of claims 1-9, wherein, The method further includes step (4): separating the product of step (3) to obtain gas, gasoline and diesel fractions, vacuum distillate oil and vacuum residue oil; wherein, the vacuum residue oil is returned to the fluidized bed reactor for recycling reaction; Preferably, in step (4), the separation is carried out by a hot high-pressure separator, a hot low-pressure separator, a cold high-pressure separator, and a cold low-pressure separator; Preferably, the hot high-pressure separator has a separation temperature of 300-450℃, a separation pressure of 8-25MPa, and a material residence time of 1-60min; Preferably, the hot low-pressure separator has a separation temperature of 250-430℃, a separation pressure of 0.01-3MPa, and a material residence time of 1-90min; Preferably, the cold high-pressure separator has a separation temperature of 20-120℃, a separation pressure of 8-25MPa, and a material residence time of 1-60min; Preferably, the cold low-pressure separator has a separation temperature of 20-120℃, a separation pressure of 0.01-3MPa, and a material residence time of 1-90min; Preferably, the ratio of the vacuum residue oil returned to the fluidized bed reactor to the fresh feedstock in step (1) is 10-20 by weight.

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