A two-stage hydrocracking process

Abstract

The present invention discloses a two-stage hydrocracking method. In the method, a graphite-containing hydrocracking catalyst is loaded into the second hydrocracking reaction zone of the second stage. The graphite content is 2% to 10% based on the mass of the graphite-containing hydrocracking catalyst, and the loading volume of the graphite-containing hydrocracking catalyst accounts for 30% to 80% of the total loading volume of all catalysts. The startup process includes: after the catalyst loading and the establishment of the inert gas circulation system are completed, the temperature of each catalyst bed is raised to 420-500°C under an inert gas environment. When the total concentration of CO and CO2 in the circulating gas reaches 0.5% to 3.0% by volume, the inlet temperature of each reactor is reduced to 140-180°C at a rate of ≯25°C / h. The two-stage hydrocracking unit is then subjected to hydrogen replacement and sulfurization. After sulfurization is completed, the feedstock is switched to introduce crude oil and normal production is resumed. This method can maintain good activity and stability of each catalyst and extend the operating cycle of the unit.

Description

Technical Field

[0001] The invention belongs to the technical field of hydrocracking, and in particular relates to a two-stage hydrocracking method. Background Art

[0002] Market demand for fuel is declining, and the proportion of fuel products in the refinery product mix is decreasing year by year. This necessitates converting a significant portion of crude oil into chemical feedstocks, which are in greater demand. The inevitable shift in refinery product mix from fuel to chemical production is imperative, requiring hydrocracking units to provide more high-quality chemical feedstocks, such as catalytic reforming feedstock and steam cracking feedstock for ethylene production. As hydrocracking units scale up, the development of two-stage hydrocracking technology for maximum production of chemical feedstocks is essential, effectively improving hydrogen utilization efficiency. While processing some feedstocks, hydrocracking units must also adjust their product mix to maximize the production of chemical feedstocks. This reduces fuel oil production while providing high-quality feedstock for catalytic reforming and ethylene units, thereby meeting the rapidly growing market demand for basic chemical feedstocks such as aromatics and ethylene.

[0003] Because hydrocracking optimizes product structure and is a powerful tool for producing chemical raw materials, it is necessary to conduct research on catalysts and processes that can maximize the production of high-quality chemical raw materials. However, in two-stage hydrocracking technology, the low hydrogen sulfide concentration in the second-stage hydrocracking reactor can easily lead to catalyst desulfurization, shortening the unit's operating cycle.

[0004] There are many reports on two-stage hydrocracking processes, but there are fewer reports on extending the operating cycle of two-stage hydrocracking units. CN112725023A discloses a two-stage hydrocracking process, in which the feedstock is subjected to hydrodesulfurization, denitrogenation, and dearomatization before entering the hydrocracking reaction zone; the products of the first hydrocracking reaction zone are separated to produce naphtha, aviation kerosene, diesel, and unconverted oil products, while the diesel and unconverted oil enter the second hydrocracking reaction zone for further cracking and then enter the fractionation system, ultimately producing naphtha, aviation kerosene products, diesel, and unconverted oil. This process can maximize the production of aviation kerosene, improve the yield of hydrocracking liquid and the yield of the target product. However, this method does not take into account the problem of low hydrogen sulfide concentration in the second-stage reactor, which can easily cause catalyst deactivation and secondary cracking.

[0005] CN110317638A discloses a two-stage heavy oil hydrocracking process. This process involves mixing the heavy oil feedstock, catalyst, and sulfiding agent with hydrogen before reacting in a first-stage hydrocracking reactor. The catalyst, desolidified / recycled tail oil, and sulfiding agent are then uniformly mixed with the effluent from the first-stage hydrocracking reactor before reacting in a second-stage hydrocracking reactor. This process can reduce the coking tendency of the heavy oil and extend the life of the catalyst in the second-stage hydrocracking reactor. However, this process requires the addition of a large amount of sulfiding agent, resulting in high costs and poor economic benefits. Summary of the Invention

[0006] In response to the shortcomings of the prior art, the present invention provides a two-stage hydrocracking method that can maintain good activity and stability of each catalyst and extend the operating cycle of the device.

[0007] The present invention provides a two-stage hydrocracking method, wherein the first stage includes a hydrofining reaction zone and a first hydrocracking reaction zone, the second stage includes a second hydrocracking reaction zone, the second hydrocracking reaction zone is filled with a hydrocracking catalyst containing graphite, the content of graphite is 2% to 10%, preferably 2% to 6%, based on the mass of the hydrocracking catalyst containing graphite, the filling volume of the hydrocracking catalyst containing graphite accounts for 30% to 80%, preferably 40% to 70% of the total filling volume of all catalysts, and the process of starting is as follows: The process includes: after completing the catalyst loading and the establishment of the inert gas circulation system in sequence, raising the temperature of each catalyst bed to 420-500°C under an inert gas environment; when it is detected that the total concentration of CO and CO2 in the circulating gas reaches 0.5v%-3.0v%, reducing the inlet temperature of each reactor to 140-180°C at a rate of ≯25°C / h (preferably 5-25°C / h); then hydrogen replacement and sulfurization are carried out on the two-stage hydrocracking unit; after the sulfurization is completed, the feed is switched to introduce crude oil and normal production is resumed.

[0008] Furthermore, a graphite-containing hydrocracking catalyst comprises an active metal component and a support, wherein the active metal component is a Group VIII or Group VIB metal, and the support comprises an acidic cracking component (such as a molecular sieve), alumina, and graphite, and may also include a binder component. The molecular sieve may be at least one of Y, Beta, and other molecular sieves. The Group VIII metal is preferably Co and / or Ni, and the Group VIB metal is preferably W and / or Mo. The graphite-containing hydrocracking catalyst is preferably prepared by the following method: uniformly mixing the acidic cracking component, alumina, water, and graphite, adding a binder, thoroughly rolling and forming the mixture, and drying to obtain a catalyst support; impregnating the catalyst support with a solution containing the active metal component, and drying the mixture in an air environment after impregnation to obtain a hydrocracking catalyst. The impregnation method may include equal volume impregnation, excess volume impregnation, and stepwise impregnation. The binder may be a conventional binder, such as at least one of glycerol, peptizing acid, octanol, citric acid, butanediol, and hexanol. The graphite is one of natural flake graphite, earthy graphite, and artificial graphite, with a particle size of 3 to 38 μm. The drying conditions for preparing the carrier are as follows: a drying temperature of 40 to 100°C, preferably 50 to 80°C. The drying time is generally 2 to 24 hours, preferably 4 to 12 hours. The drying conditions after impregnation are as follows: a drying temperature of 100 to 200°C, preferably 120 to 150°C, and a drying time of 2 to 24 hours, preferably 4 to 12 hours. Based on the mass of the hydrocracking catalyst containing the graphite, the carrier content is 65% to 90%, and the active metal content, calculated as oxide, is 10% to 35%. Based on the mass of the hydrocracking catalyst containing the graphite, the acidic cracking component content is 10% to 40%, the alumina content is 30% to 55%, the graphite content is 2% to 10%, and the binder content is less than 10%. Based on the mass of the hydrocracking catalyst containing graphite, the content of the Group VIII metal in terms of oxide is 1% to 8%, and the content of the Group VIB metal in terms of oxide is 15% to 30%.

[0009] Furthermore, the first hydrorefining reaction zone is loaded with a hydrorefining catalyst, and the first hydrocracking reaction zone is loaded with a first hydrocracking catalyst. These catalysts can be prepared according to existing methods, or industrial catalysts can be used. For example, the hydrorefining catalyst comprises a hydrogenation-active metal component, a support, and a binder. The hydrogenation-active metal comprises at least one of a Group VIB metal (such as tungsten and / or molybdenum) and a Group VIII metal (such as nickel and / or cobalt); the support comprises one or more of alumina, silica, and amorphous silica-alumina. Based on the weight of the catalyst, the Group VIB metal content, calculated as oxide, is 15% to 40%, preferably 25% to 30%, and the Group VIII metal content, calculated as oxide, is 2% to 10%, preferably 3% to 8%. Hydrorefining catalysts such as FF-26, FF-56, FF-76A, and FF-76 developed by the Dalian Research Institute of Petrochemicals can be used. The first hydrocracking catalyst comprises a hydrogenation-active metal component, a cracking component, and a binder. The hydrogenation-active metal comprises at least one of a Group VIB metal (such as tungsten and / or molybdenum) and a Group VIII metal (such as nickel and / or cobalt). The cracking component comprises amorphous silica-alumina and / or a molecular sieve, wherein the molecular sieve comprises at least one of a Y-type, a β-type, or a ZSM-5 type molecular sieve. Based on the weight of the catalyst, the Group VIB metal content, calculated as oxide, is 15% to 30%, preferably 18% to 27%. The Group VIII metal content, calculated as oxide, is 1% to 8%, preferably 2% to 6%. The molecular sieve content is 10% to 45%. The first hydrocracking catalyst may be selected from existing hydrocracking catalysts such as FC-14, FC-16B, FC-32, FC-32A, FC-86, or FC-52 developed by the Dalian Petrochemical Research Institute. In the second hydrocracking reaction zone in the second stage, after the hydrocracking catalyst, a post-hydrogenation refining catalyst can be loaded. The catalyst can be the same as the above-mentioned hydrofining catalyst, or FF-12 developed by Dalian Petrochemical Research Institute can be used.

[0010] Furthermore, the hydrogen replacement is completed when the hydrogen volume fraction in the circulating hydrogen reaches 95% to 99%. The hydrogen replacement process can introduce new hydrogen into the circulating gas, and introduce it into the hydrocracking unit at the inlet of the hydrofining reaction zone and the inlet of the second hydrocracking reaction zone to replace the gas in the system, until the hydrogen volume fraction in the circulating hydrogen reaches 95% to 99%, at which time the hydrogen replacement is completed.

[0011] Furthermore, when new hydrogen is introduced into the circulating hydrogen during the hydrogen replacement process, the inlet temperature of the second hydrocracking reaction zone needs to be lowered to the same temperature as the inlet temperature of the hydrofining reaction zone.

[0012] Furthermore, the vulcanization process preferably adopts wet vulcanization.

[0013] Furthermore, the wet sulfurization includes a first sulfurization stage and a second sulfurization stage, wherein the first sulfurization stage is to increase the inlet temperature of the hydrofining reaction zone and the second hydrocracking reaction zone to a first constant temperature of 230-250°C at a heating rate of ≯10°C / h, preferably 5-10°C / h, and then maintain the temperature at the first constant temperature for 4-12 hours. In the first sulfurization stage, the mass concentration of hydrogen sulfide in the circulating hydrogen is controlled to be maintained at 1000-5000 ppm; and the second sulfurization stage is to increase the inlet temperature of the hydrofining reaction zone and the second hydrocracking reaction zone to a second constant temperature of 330-370°C, preferably 340-370°C, at a heating rate of ≯10°C / h, preferably 3-10°C / h, and then maintain the temperature at the second constant temperature for 4-12 hours. In the second sulfurization stage, the mass concentration of hydrogen sulfide in the circulating hydrogen is controlled to reach 5000-10000 ppm.

[0014] Furthermore, the wet sulfurization process allows hydrogen sulfide to penetrate the catalyst bed before entering the first sulfurization stage. The process is as follows: low-nitrogen oil is introduced from the inlet of the hydrorefining reaction zone to wet the catalyst, a closed-loop circulation of the low-nitrogen oil is established, and the inlet temperature of the hydrorefining reaction zone and the second hydrocracking reaction zone is increased to 190-210°C at a rate of ≯15°C / h (preferably 5-15°C / h), and then a sulfurizing agent is gradually injected into the low-nitrogen oil until hydrogen sulfide penetrates the entire catalyst bed. Preferably, the conditions for hydrogen sulfide to penetrate the entire catalyst bed are as follows: the mass concentration of hydrogen sulfide in the circulating hydrogen is above 1000 ppm, more preferably 1000-5000 ppm.

[0015] After the first constant-temperature sulfidation process is completed, ammonia injection passivation is initiated. The sulfiding agent injection rate is adjusted to maintain a hydrogen sulfide concentration of 5,000 to 10,000 ppm in the circulating hydrogen. The inlet temperature of the two reactors is steadily increased at a rate of ≤8°C / h. Before ammonia breakthrough, the maximum catalyst bed temperature in the first and second hydrocracking reaction zones is controlled to ≤260°C. Ammonia breakthrough is considered to have occurred when the ammonia content in the high-density acidic water exceeds 1,000 ppm. After ammonia breakthrough, the passivating agent injection rate is adjusted to maintain a ammonia concentration in the high-density acidic water ≤5,000 ppm, preferably between 5,000 and 15,000 ppm. Passivating agent addition is discontinued until the constant-temperature sulfidation process is completed.

[0016] Furthermore, the condition for ammonia breakthrough is that the mass concentration of ammonium ions in the acidic water discharged from the bottom of the high-pressure separator reaches 1000 ppm or more, preferably 1000 to 20000 ppm.

[0017] Furthermore, after the passivation is completed, the feed is switched. After all the raw oil is introduced, the temperature of the hydrorefining reaction zone and the second hydrocracking reaction zone is adjusted until the yield and quality of the target product are qualified, so that normal production can be carried out.

[0018] Furthermore, the vulcanizing agent is an industrially commonly used vulcanizing agent and / or an environmentally friendly vulcanizing agent, such as at least one of dimethyl disulfide (DMDS) and carbon disulfide (CS2).

[0019] Furthermore, the passivating agent is selected from one or more of anhydrous liquid ammonia, cyclohexylamine, aniline, and n-butylamine.

[0020] Furthermore, when the passivating agent is selected from anhydrous liquid ammonia, the passivating agent injection port is set on the high-pressure pipeline between the outlet of the feed oil pump and the reaction inlet, and a high-pressure ammonia injection pump is required. When the passivating agent is selected from a passivating agent other than anhydrous liquid ammonia, the passivating agent can be directly injected into the feed oil tank or into the oil inlet pipeline from the feed oil tank to the feed oil pump.

[0021] Furthermore, the low-nitrogen oil is generally atmospheric distillate oil, with an initial distillation point generally between 150 and 230° C., a final distillation point generally ≯360° C., a nitrogen mass content ≯300 ppm, and a water mass content ≯100 ppm.

[0022] Furthermore, the hydrofining reaction zone and the first hydrocracking reaction zone may be one reactor and / or multiple reactors.

[0023] Furthermore, the two-stage hydrocracking method can produce at least one target product among heavy naphtha, diesel, kerosene, etc. according to actual needs; the production of heavy naphtha is now described as an example:

[0024] The feedstock oil is mixed with hydrogen and sequentially enters the first hydrorefining reaction zone and the first hydrocracking reaction zone. The resulting reaction effluent is fractionated to produce gas, light naphtha, heavy naphtha products, and a tail oil fraction. The tail oil fraction and hydrogen are then reacted in the second hydrocracking reaction zone of the second stage. The resulting second hydrocracking reaction zone effluent enters the fractionation system (preferably, the second hydrocracking reaction zone effluent and the first hydrocracking reaction zone effluent share the same fractionation system) to produce gas, light naphtha, heavy naphtha products, and a tail oil fraction. The initial boiling point of the tail oil fraction is generally 160-370°C.

[0025] Furthermore, the raw oil is selected from wax oil and / or diesel raw materials, the initial boiling point of the raw materials is generally 155-380°C, the final boiling point is generally 390-750°C, the sulfur content is 0.2wt%-3.0wt%, and the nitrogen mass content is 500ppm-2000ppm. The wax oil raw material can be selected from one or more of vacuum wax oil, coker wax oil, ebullated bed wax oil and slurry bed wax oil. The diesel raw material can be selected from one or more of straight-run diesel, coker diesel, catalytic diesel, ebullated bed diesel and slurry bed diesel. When processing inferior raw oil, a protective agent is preferably loaded on the top of the hydrorefining reaction zone, which can be the FZC series and FBN protective agents developed by Dalian Petrochemical Research Institute.

[0026] Furthermore, during normal production, the reaction conditions in the first stage hydrofining reaction zone and the first cracking reaction zone are as follows:

[0027] The reaction pressure is 6.0-20.0 MPa, preferably 9.0-16.0 MPa; the average reaction temperature is 270-440°C, preferably 320-415°C; the hydrogen-oil volume ratio is 400:1-2200:1, preferably 600:1-1400:1; the hydrorefining volume space velocity is 0.2-5.0 h -1 , preferably 0.5~4.0h -1 ; Hydrocracking volume space velocity is 0.8~5.0h -1 , preferably 0.5~3.0h -1 .

[0028] Furthermore, during normal production, the reaction conditions of the second hydrocracking reaction zone in the second stage are as follows:

[0029] The reaction pressure is 6.0-20.0 MPa, preferably 9.0-16.0 MPa; the average reaction temperature is 230-400°C, preferably 300-360°C; the total volume space velocity is 0.5-4.0 h -1 , preferably 1.0~4.0h -1 ; The volume ratio of hydrogen to oil is 400:1 to 2200:1, preferably 800:1 to 1600:1.

[0030] Compared with the prior art, the method of the present invention has the following beneficial effects:

[0031] 1. The active metals in the oxidation state on the surface of the hydrogenation catalyst are divided into hexacoordinated and tetracoordinated molybdenum or tungsten species. The hexacoordinated species H2-TPR has a low reduction peak temperature and good hydrogenation activity, but its stability is poor and its resistance to desulfurization is weak; while the tetracoordinated species H2-TPR has a high reduction peak temperature, good stability and strong resistance to desulfurization. In the two-stage hydrocracking process, the concentration of hydrogen sulfide in the second-stage hydrogenation reactor is low, and the catalyst is in a desulfurized state for a long time. Therefore, it is necessary to load a high-stability hydrogenation catalyst to ensure long-term operation of the device. The present invention loads a graphite-containing hydrocracking catalyst into the second hydrocracking reaction zone of the second stage. When the device is started, the appropriate reduction conditions are controlled under an inert gas atmosphere to moderately convert the hexacoordinated species in the catalyst into a tetracoordinated molybdenum species, and adjust the sulfurization conditions, especially increase the secondary constant-temperature sulfurization temperature, thereby improving the stability of the catalyst while ensuring the activity of the catalyst.

[0032] 2. When the two-stage hydrocracking method of the present invention is used for two-stage hydrocracking production of heavy naphtha, the second hydrocracking catalyst is highly stable, has good hydrogenation performance, reduces hydrogen consumption, and can significantly increase the aromatic potential of heavy naphtha. When used as a reforming feedstock, the hydrogen yield and product quality are significantly improved. DETAILED DESCRIPTION

[0033] The two-stage hydrocracking process provided by the present invention will be further described below with reference to the examples, but the present invention is not limited thereto.

[0034] Table 1 Properties of low nitrogen start-up oil and feed oil

[0035]

[0036]

[0037] Table 2 Industrial catalysts

[0038] Industrial agents FF-12 FF-76 FC-86 FC-46 Physical and chemical properties Average pore size / nm 6.0nm 5.6nm 4.5nm 4.2nm <![CDATA[Pore volume / mL·g -1 > 0.30 0.35 0.28 0.26 <![CDATA[Specific surface area / m 2 ·g -1 > 165 192 310 305 shape Shamrock Gear Ball Gear Ball Gear Ball <![CDATA[Loading heap ratio, g / cm 3 > 0.64 0.75 0.68 0.81

[0039] Table 3 Evaluation conditions

[0040] High pressure, MPa 14.0 <![CDATA[Volume hourly space velocity of the refining agent in the first-stage refining reactor, h -1 > 1.2 <![CDATA[The hourly space velocity of the cracking agent in a cracking reactor -1 > 1.4 <![CDATA[Volume hourly space velocity of the cracking agent in the two-stage cracking reactor, h -1 > 1.8 First stage / second stage cracking reaction>350℃ conversion rate, % 65 / 50 Hydrogen-to-oil ratio at the inlet of the first-stage hydrotreating / second-stage hydrocracking reaction zone 900:1 / 1100:1 Nitrogen content of refined oil, ppm ～10 Running time, h 2000

[0041] The two-stage hydrocracking process used in the following examples and comparative examples employs two reactors, wherein the first reactor comprises a first hydrorefining reaction zone (loaded with a hydrorefining catalyst) and a first hydrocracking reaction zone (loaded with a hydrocracking catalyst), and the second reactor comprises a second hydrocracking reaction zone (loaded with a hydrocracking catalyst and a post-refining catalyst), and comprises: a feedstock oil and hydrogen are mixed and sequentially enter the first hydrorefining reaction zone and the first hydrocracking reaction zone, and the resulting reaction effluent enters a fractionation system to obtain gas, light naphtha, heavy naphtha products, and a tail oil fraction; the tail oil fraction and hydrogen enter the second hydrocracking reaction zone of the second stage, and the resulting second hydrocracking reaction zone effluent enters the above-mentioned fractionation system, i.e., the effluent of the second hydrocracking reaction zone and the effluent of the first hydrocracking reaction zone share a fractionation system.

[0042] The properties of the start-up low nitrogen oil and feed oil used in the following examples and comparative examples are shown in Table 1. In all examples and comparative examples, the first refining reaction zone was filled with FF-76 hydrorefining catalyst, the first cracking reactor zone was filled with FC-46 hydrocracking catalyst, and the second hydrocracking reaction zone was filled with the hydrocracking catalyst provided in the examples or comparative examples. The bottom was filled with liquid at a volume space velocity of 20.0 h -1 The distillation range of light naphtha in the examples and comparative examples is <65°C, the distillation range of heavy naphtha is 65-175°C, and the distillation range of tail oil is >175°C.

[0043] Example 1

[0044] The preparation method of the hydrocracking catalyst containing graphite loaded in the second-stage hydrocracking reactor is as follows:

[0045] A catalyst support was synthesized by doping a mixture of Y molecular sieve and alumina with graphite (average particle size 10 μm). Water and a binder of glycerol were then added. The mixture was extruded and dried at 80°C for 4 hours before being extruded. The active metals were impregnated with nickel nitrate hexahydrate as the nickel source and ammonium molybdate hexahydrate as the molybdenum source using an equal-volume co-impregnation method. After drying at 120°C for 4 hours, a graphite-containing hydrocracking catalyst, designated Cat-1, was obtained. The catalyst composition, based on catalyst mass, was 25% Y molecular sieve, 40% alumina, 3% binder, and 3% graphite. The metal content, calculated as nickel oxide, was 3.5% by weight, and the metal content, calculated as molybdenum oxide, was 25.5% by weight. The resulting Cat-1 catalyst was loaded into the second hydrocracking reaction zone of a two-stage hydrocracking unit.

[0046] After catalyst loading and nitrogen recirculation system establishment in the two-stage hydrocracking unit, the inlet temperature of the first-stage refining reactor was raised to 430°C, and the inlet temperature of the second-stage cracking reactor was raised to 430°C. These temperatures were maintained for a period of time. When the CO+CO2 volume fraction in the circulating gas exceeded 0.5% by volume, the inlet temperature of each reactor was reduced at a rate of 20°C / h to 160°C, in preparation for the introduction of fresh hydrogen and a gradual increase in the reaction system pressure. Hydrogen was introduced into the hydrocracking unit from the inlet of the hydrofining reaction zone and the inlet of the second cracking reaction zone, replacing the gas in the reaction system. When the volume fraction of hydrogen in the circulating hydrogen reached 96%, hydrogen replacement was completed, and the system pressure was subsequently increased.

[0047] When the system pressure increases to 14.0MPa, prepare to introduce low-nitrogen start-up oil. After the oil is introduced, increase the inlet temperature of each reactor to 210°C at a rate of 12°C / h, then gradually inject sulfiding agent into the low-nitrogen oil. When the hydrogen sulfide concentration in the circulating hydrogen is >1000ppm, begin to increase the inlet temperature of each reactor to 240°C at a rate of 8°C / h. After the temperature is reached, begin constant temperature sulfidation for 8 hours, during which the hydrogen sulfide concentration is maintained at 1000-5000ppm. After the constant temperature sulfidation is completed, begin liquid ammonia passivation and adjust the injection rate of the sulfiding agent to ensure that the hydrogen sulfide concentration in the circulating hydrogen reaches 5000-10000ppm. Wait for ammonia breakthrough at 250°C. When the ammonium ion concentration in the acidic water at the bottom of the high-precision reactor reaches above 1000ppm, continue to steadily increase the inlet temperature of the two-stage reactor to 340°C at a rate of 6°C / h and sulfidate at constant temperature for 6 hours. After the constant temperature test, we prepared to switch to Iranian VGO feedstock and adjusted the refining and cracking reaction temperatures until the product yield and quality were acceptable. Subsequently, we conducted a 2000-hour long-term stability evaluation experiment according to the process conditions in Table 3.

[0048] Example 2

[0049] The preparation method of the graphite catalyst loaded in the second-stage hydrocracking reactor is as follows:

[0050] A catalyst support was synthesized by doping a mixture of Y molecular sieve and alumina with graphite (average particle size 18 μm). Water and a binder of glycerol were then added. The mixture was extruded and dried at 80°C for 4 hours before being extruded. The active metals were impregnated with nickel nitrate hexahydrate as the nickel source and ammonium molybdate hexahydrate as the molybdenum source using an equal-volume co-impregnation method. After drying at 120°C for 4 hours, a graphite-containing hydrocracking catalyst, designated Cat-2, was obtained. The catalyst weight was used as the basis: 35% Y molecular sieve, 30% alumina, 4% binder, and 2% graphite. The metal content, calculated as nickel oxide, was 4.5% by weight, and the metal content, calculated as molybdenum oxide, was 24.5% by weight. The resulting Cat-2 catalyst was loaded into the second-stage hydrocracking reactor of a two-stage hydrocracking unit.

[0051] After catalyst loading and nitrogen recirculation system establishment in the two-stage hydrocracking unit, the inlet temperature of the first-stage refining reactor was raised to 445°C, and the inlet temperature of the second-stage cracking reactor was raised to 445°C. These temperatures were maintained for a period of time. When the CO+CO2 volume fraction in the circulating gas exceeded 1.0% by volume, the inlet temperature of each reactor was reduced at a rate of 20°C / h to 160°C, in preparation for the introduction of fresh hydrogen and a gradual increase in the reaction system pressure. Hydrogen was introduced into the hydrocracking unit from the inlet of the hydrofining reaction zone and the inlet of the second cracking reaction zone, replacing the gas in the reaction system. When the volume fraction of hydrogen in the circulating hydrogen reached 97%, hydrogen replacement was completed, and the system pressure was subsequently increased.

[0052] When the system pressure increases to 14.0 MPa, prepare to introduce low-nitrogen start-up oil. After the oil is introduced, increase the inlet temperature of each reactor to 220°C at a rate of 10°C / h, then gradually inject sulfiding agent into the low-nitrogen oil. When the hydrogen sulfide concentration in the circulating hydrogen is >1000 ppm, begin to increase the inlet temperature of each reactor to 235°C at a rate of 8°C / h. Once the temperature is reached, begin constant temperature sulfidation for 8 hours, during which the hydrogen sulfide concentration is maintained at 1000-5000 ppm. After the constant temperature sulfidation is completed, begin liquid ammonia passivation and adjust the sulfiding agent injection rate to ensure that the hydrogen sulfide concentration in the circulating hydrogen reaches 5000-10000 ppm. Wait for ammonia breakthrough at 255°C. When the ammonium ion concentration in the acidic water at the bottom of the high-precision reactor reaches above 2000 ppm, continue to steadily increase the inlet temperature of the two-stage reactor to 350°C at a rate of 6°C / h and sulfidate at constant temperature for 6 hours. After the constant temperature test, we prepared to switch to Iranian VGO feedstock and adjusted the refining and cracking reaction temperatures until the product yield and quality were acceptable. Subsequently, we conducted a 2000-hour long-term stability evaluation experiment according to the process conditions in Table 3.

[0053] Example 3

[0054] The preparation method of the graphite catalyst loaded in the second-stage hydrocracking reactor is as follows:

[0055] A catalyst support was synthesized by doping a mixture of Y molecular sieve and alumina with graphite (average particle size 25 μm). Water and a binder of glycerol were then added, extruded, and dried at 90°C for 4 hours. The active metals were impregnated with nickel nitrate hexahydrate as the nickel source and ammonium molybdate hexahydrate as the molybdenum source using an equal-volume co-impregnation method. After drying at 120°C for 4 hours, a graphite-containing hydrocracking catalyst, designated Cat-3, was obtained. The catalyst weight was used as the basis: 30 wt% Y molecular sieve, 35 wt% alumina, 3 wt% binder, and 3 wt% graphite. The metal content, calculated as nickel oxide, was 4.0 wt%, and the metal content, calculated as molybdenum oxide, was 25.0 wt%. The resulting Cat-3 catalyst was loaded into the second-stage hydrocracking reactor of a two-stage hydrocracking unit.

[0056] After catalyst loading and nitrogen recirculation system establishment in the two-stage hydrocracking unit, the inlet temperature of the first-stage refining reactor was raised to 450°C, and the inlet temperature of the second-stage cracking reactor was raised to 450°C. These temperatures were maintained for a period of time. When the CO+CO2 volume fraction in the circulating hydrogen exceeded 1.5% by volume, the inlet temperature of each reactor was reduced at a rate of 20°C / h to 160°C, in preparation for the introduction of fresh hydrogen and a gradual increase in the reaction system pressure. Hydrogen was introduced into the hydrocracking unit from the inlet of the hydrofining reaction zone and the inlet of the second cracking reaction zone, replacing the gas in the reaction system. When the volume fraction of hydrogen in the circulating hydrogen reached 98%, hydrogen replacement was completed, and the system pressure was subsequently increased.

[0057] When the system pressure increases to 14.0MPa, prepare to introduce low-nitrogen start-up oil. After the oil is introduced, increase the inlet temperature of each reactor to 210°C at a rate of 10°C / h, then gradually inject sulfiding agent into the low-nitrogen oil. When the hydrogen sulfide concentration in the circulating hydrogen is >1000ppm, begin to increase the inlet temperature of each reactor to 245°C at a rate of 8°C / h. After the temperature is reached, begin constant temperature sulfidation for 8 hours, during which the hydrogen sulfide concentration is maintained at 1000-5000ppm. After the constant temperature sulfidation is completed, begin liquid ammonia passivation and adjust the injection rate of the sulfiding agent to ensure that the hydrogen sulfide concentration in the circulating hydrogen reaches 5000-10000ppm. Wait for ammonia breakthrough at 250°C. When the ammonium ion concentration in the acidic water at the bottom of the high-precision reactor reaches above 3000ppm, continue to steadily increase the inlet temperature of the two-stage reactor to 360°C at a rate of 6°C / h and sulfidate at constant temperature for 6 hours. After the constant temperature test, we prepared to switch to Iranian VGO feedstock and adjusted the refining and cracking reaction temperatures until the product yield and quality were acceptable. Subsequently, we conducted a 2000-hour long-term stability evaluation experiment according to the process conditions in Table 3.

[0058] Example 4

[0059] The preparation method of the graphite catalyst loaded in the second-stage hydrocracking reactor is as follows:

[0060] A catalyst support was synthesized by doping a mixture of Y molecular sieve and alumina with graphite (average particle size 30 μm). Water and a binder of glycerol were then added. The catalyst was extruded and dried at 80°C for 4 hours. The active metals were impregnated with nickel nitrate hexahydrate as the nickel source and ammonium molybdate hexahydrate as the molybdenum source using an equal-volume co-impregnation method. After drying at 140°C for 4 hours, the resulting graphite-containing hydrocracking catalyst, designated Cat-4, contained 25% Y molecular sieve, 40% alumina, 2% binder, and 4% graphite, based on the catalyst mass. The metal content, calculated as nickel oxide, was 4.5% by weight, and the metal content, calculated as molybdenum oxide, was 24.5% by weight. The resulting Cat-4 catalyst was loaded into the second-stage hydrocracking reactor of a two-stage hydrocracking unit.

[0061] After catalyst loading and nitrogen recirculation system establishment in the two-stage hydrocracking unit, the inlet temperature of the first-stage refining reactor was raised to 460°C, and the inlet temperature of the second-stage cracking reactor was raised to 460°C. These temperatures were maintained for a period of time. When the CO+CO2 volume fraction in the circulating hydrogen exceeded 2.5% by volume, the inlet temperature of each reactor was reduced at a rate of 15°C / h to 160°C, in preparation for the introduction of fresh hydrogen and a gradual increase in the reaction system pressure. Hydrogen was introduced into the hydrocracking unit through the inlet of the hydrofining reaction zone and the inlet of the second cracking reaction zone, replacing the gas in the reaction system. When the volume fraction of hydrogen in the circulating hydrogen reached 99%, hydrogen replacement was completed, and the system pressure was subsequently increased.

[0062] When the system pressure increases to 14.0MPa, prepare to introduce low-nitrogen start-up oil. After the oil is introduced, increase the reactor inlet temperature to 215°C at a rate of 10°C / h, then gradually inject sulfiding agent into the low-nitrogen oil. When the hydrogen sulfide concentration in the circulating hydrogen is >1000ppm, begin to increase the reactor inlet temperature to 240°C at a rate of 8°C / h. Once the temperature is reached, begin constant-temperature sulfidation for 8 hours, during which the hydrogen sulfide concentration is maintained at 1000-5000ppm. After the constant-temperature sulfidation is completed, begin liquid ammonia passivation and adjust the sulfiding agent injection rate to ensure that the hydrogen sulfide concentration in the circulating hydrogen reaches 5000-10000ppm. Wait for ammonia breakthrough at 260°C. When the ammonium ion concentration in the acidic water at the bottom of the high-precision reactor reaches above 4000ppm, continue to steadily increase the inlet temperature of the two-stage reactor to 370°C at a rate of 6°C / h and sulfidate at constant temperature for 6 hours. After the constant temperature test, we prepared to switch to Iranian VGO feedstock and adjusted the refining and cracking reaction temperatures until the product yield and quality were acceptable. Subsequently, we conducted a 2000-hour long-term stability evaluation experiment according to the process conditions in Table 3.

[0063] Example 5

[0064] The preparation method of the graphite-containing hydrocracking catalyst loaded in the second-stage hydrocracking reactor is the same as that in Example 1.

[0065] After catalyst loading and nitrogen recirculation system establishment in the two-stage hydrocracking unit, the inlet temperature of the first-stage refining reactor was raised to 430°C, and the inlet temperature of the second-stage cracking reactor was raised to 430°C. These temperatures were maintained for a period of time. When the CO+CO2 volume fraction in the circulating gas exceeded 0.5% by volume, the inlet temperature of each reactor was reduced at a rate of 20°C / h to 160°C, in preparation for the introduction of fresh hydrogen and a gradual increase in the reaction system pressure. Hydrogen was introduced into the hydrocracking unit from the inlet of the hydrofining reaction zone and the inlet of the second cracking reaction zone, replacing the gas in the reaction system. When the volume fraction of hydrogen in the circulating hydrogen reached 96%, hydrogen replacement was completed, and the system pressure was subsequently increased.

[0066] When the system pressure increases to 14.0MPa, prepare to introduce low-nitrogen start-up oil. After the oil is introduced, increase the inlet temperature of each reactor to 210°C at a rate of 12°C / h, then gradually inject sulfiding agent into the low-nitrogen oil. When the hydrogen sulfide concentration in the circulating hydrogen is >1000ppm, begin to increase the inlet temperature of each reactor to 240°C at a rate of 8°C / h. Once the temperature is reached, begin constant temperature sulfidation for 8 hours, during which the hydrogen sulfide concentration is maintained at 1000-5000ppm. After the constant temperature sulfidation is completed, begin liquid ammonia passivation and adjust the sulfiding agent injection rate to ensure that the hydrogen sulfide concentration in the circulating hydrogen reaches 5000-10000ppm. Wait for ammonia breakthrough at 250°C. When the ammonium ion concentration in the acidic water at the bottom of the high-precision reactor reaches above 1000ppm, continue to steadily increase the inlet temperature of the two-stage reactor to 330°C at a rate of 6°C / h and continue constant temperature sulfidation for 8 hours. After the constant temperature test, we prepared to switch to Iranian VGO feedstock and adjusted the refining and cracking reaction temperatures until the product yield and quality were acceptable. Subsequently, we conducted a 2000-hour long-term stability evaluation experiment according to the process conditions in Table 3.

[0067] Comparative Example 1

[0068] The first-stage refining reactor was loaded with FF-76 hydrofining catalyst, the first-stage cracking reactor was loaded with FC-46 hydrocracking catalyst, and the second-stage cracking reactor was loaded with FC-86 hydrocracking catalyst. After sulfurization according to conventional startup procedures, a 2000-hour long-term stability evaluation experiment was conducted using Iranian VGO as the raw material, according to the process conditions in Table 3. The conventional startup sulfurization method is as follows:

[0069] After completing catalyst loading, establishing the circulating hydrogen system, and completing airtightness work in the two-stage hydrocracking unit, and when the system pressure reaches 14.0 MPa, preparations are made to introduce low-nitrogen start-up oil. After oil introduction, the reactor inlet temperature is raised to 200°C at a rate of 10°C / h, and sulfiding agent injection begins. When the hydrogen sulfide concentration in the circulating hydrogen reaches >1000 ppm, the reactor inlet temperature is raised to 230°C at a rate of 8°C / h. Once the temperature is reached, constant-temperature sulfidation begins for 8 hours, maintaining the hydrogen sulfide concentration between 1000 and 5000 ppm. After constant-temperature sulfidation is completed, liquid ammonia passivation begins, and the sulfiding agent injection rate is adjusted to ensure that the hydrogen sulfide concentration in the circulating hydrogen reaches 5000 to 10000 ppm. After ammonia breakthrough, the inlet temperature of the two-stage reactor is steadily raised to 290°C at a rate of 6°C / h, and constant-temperature sulfidation is continued for 2 hours. After the constant temperature was completed, the inlet temperature of the two reactors was further increased to 310°C in preparation for switching to Iranian VGO feedstock. The refining and cracking reaction temperatures were adjusted until the product yield and quality met the standards. Subsequently, a 2000-hour long-term stability evaluation experiment was conducted according to the process conditions in Table 3.

[0070] Comparative Example 2

[0071] The first-stage refining reactor was loaded with FF-76 hydrofining catalyst, the first-stage cracking reactor was loaded with FC-46 hydrocracking catalyst, and the second-stage cracking reactor was loaded with FC-46 hydrocracking catalyst. After sulfurization according to conventional startup procedures, a 2000-hour long-term stability evaluation experiment was conducted using Iranian VGO as the raw material, according to the process conditions in Table 3. The conventional startup sulfurization method is as follows:

[0072] After completing catalyst loading, establishing the circulating hydrogen system, and achieving airtightness in the two-stage hydrocracking unit, and when the system pressure reaches 14.0 MPa, preparations are made to introduce low-nitrogen start-up oil. Following oil introduction, the reactor inlet temperature is raised to 210°C at a rate of 10°C / h, and sulfiding agent injection begins. When the hydrogen sulfide concentration in the circulating hydrogen reaches >1000 ppm, the reactor inlet temperature is raised to 230°C at a rate of 8°C / h. Once the temperature is reached, constant-temperature sulfidation is initiated for 4 hours, maintaining the hydrogen sulfide concentration between 1000 and 5000 ppm. Following constant-temperature sulfidation, liquid ammonia passivation is initiated, and the sulfiding agent injection rate is adjusted to maintain the hydrogen sulfide concentration in the circulating hydrogen between 5000 and 10000 ppm. Once ammonia breakthrough is achieved, the inlet temperature of the two-stage reactor is steadily raised to 290°C at a rate of 6°C / h, and constant-temperature sulfidation is continued for 2 hours. After the constant temperature was completed, the inlet temperature of the two reactors was further increased to 310°C in preparation for switching to Iranian VGO feedstock. The refining and cracking reaction temperatures were adjusted until the product yield and quality met the standards. Subsequently, a 2000-hour long-term stability evaluation experiment was conducted according to the process conditions in Table 3.

[0073] Comparative Example 3

[0074] The preparation method of the conventional catalyst loaded in the second-stage hydrocracking reactor is as follows:

[0075] A catalyst support was synthesized using a mixture of Y molecular sieve and alumina as the carrier component, with glycerol added as a binder. Extrusion was performed and dried at 80°C for 4 hours before synthesizing the catalyst support. Nickel nitrate hexahydrate was used as the nickel source, and ammonium molybdate hexahydrate was used as the molybdenum source. The active metals were impregnated using an equal-volume co-impregnation method. After drying at 120°C for 4 hours, the resulting hydrocracking catalyst, designated Cat-5, contained 25% Y molecular sieve, 40% alumina, and 6% binder, based on the catalyst mass. The metal content, calculated as nickel oxide, was 3.5% by weight, and the metal content, calculated as molybdenum oxide, was 25.5% by weight. The resulting Cat-5 catalyst was loaded into the second-stage hydrocracking reactor of a two-stage hydrocracking unit.

[0076] After catalyst loading and the establishment of the circulating hydrogen system in the two-stage hydrocracking unit, when the system pressure reaches 14.0 MPa, preparations are made to introduce low-nitrogen start-up oil. Following oil introduction, the reactor inlet temperature is raised to 210°C at a rate of 12°C / h, and sulfiding agent injection begins. When the hydrogen sulfide concentration in the circulating hydrogen exceeds 1000 ppm, the reactor inlet temperature is raised to 240°C at a rate of 8°C / h. Once the temperature is reached, constant-temperature sulfiding begins for 8 hours, maintaining the hydrogen sulfide concentration between 1000 and 5000 ppm. Following constant-temperature sulfiding, liquid ammonia passivation is initiated, and the sulfiding agent injection rate is adjusted to maintain the hydrogen sulfide concentration in the circulating hydrogen between 5000 and 10000 ppm. Ammonia breakthrough is awaited at 250°C. When the ammonium ion concentration in the acidic water at the high-fraction bottom reaches above 1000 ppm, the inlet temperature of the two-stage reactor is steadily raised to 340°C at a rate of 6°C / h, and constant-temperature sulfiding continues for 6 hours. After the constant temperature test, we prepared to switch to Iranian VGO feedstock and adjusted the refining and cracking reaction temperatures until the product yield and quality were acceptable. Subsequently, we conducted a 2000-hour long-term stability evaluation experiment according to the process conditions in Table 3.

[0077] Table 4 Test results of each embodiment

[0078]

[0079] Table 5 Test results of each comparative example

[0080]

[0081] It can be seen from the experimental results of the comparative examples and the embodiments that, under the same two-stage conversion rate, the two-stage hydrocracking process of the present invention not only produces high heavy naphtha yield and aromatic potential, but also has low hydrogen consumption and slow catalyst deactivation rate.

Claims

1. A two-stage hydrocracking process, wherein: The first stage includes a hydrorefining reaction zone and a first hydrocracking reaction zone, and the second stage includes a second hydrocracking reaction zone. The second hydrocracking reaction zone is loaded with a hydrocracking catalyst containing graphite. The graphite content is 2% to 10% based on the mass of the hydrocracking catalyst containing graphite, and the loading volume of the hydrocracking catalyst containing graphite accounts for 30% to 80% of the total loading volume of all catalysts. The start-up process includes: after the catalyst loading is completed and the inert gas circulation system is established, the temperature of each catalyst bed is increased to 420 to 500°C under an inert gas environment. When the total concentration of CO and CO2 in the circulating gas is detected to reach 0.5v% to 3.0v%, the inlet temperature of each reactor is reduced to 140 to 180°C at a rate of ≯25°C / h. Then, the two-stage hydrocracking units are subjected to hydrogen replacement and sulfurization. After the sulfurization is completed, the feed is switched to introduce crude oil and normal production is resumed. The graphite-containing hydrocracking catalyst comprises an active metal component and a carrier, wherein the active metal component is a Group VIII or Group VIB metal, and the carrier comprises an acidic cracking component, alumina and graphite; The sulfiding process adopts wet sulfiding, and the wet sulfiding includes a first sulfiding stage and a second sulfiding stage. In the first sulfiding stage, the inlet temperatures of the hydrofining reaction zone and the second hydrocracking reaction zone are increased to a first constant temperature of 230-250° C. at a heating rate of ≯10° C. / h, and then maintained at the first constant temperature for 4-12 hours. In the first sulfiding stage, the mass concentration of hydrogen sulfide in the circulating hydrogen is controlled to be maintained at 1000-5000 ppm; in the second sulfiding stage, the inlet temperatures of the hydrofining reaction zone and the second hydrocracking reaction zone are increased to a second constant temperature of 330-370° C. at a heating rate of ≯10° C. / h, and then maintained at the second constant temperature for 4-12 hours. In the second sulfiding stage, the mass concentration of hydrogen sulfide in the circulating hydrogen is controlled to reach 5000-10000 ppm.

2. The method according to claim 1, characterized in that In the hydrocracking catalyst containing graphite, the acidic cracking component is a molecular sieve selected from at least one of Y and Beta molecular sieves; the Group VIII metal is Co and / or Ni, and the Group VIB metal is W and / or Mo.

3. The method according to claim 1, characterized in that Based on the mass of the hydrocracking catalyst containing graphite, the carrier content is 65% to 90%, and the content of the active metal in terms of oxide is 10% to 35%.

4. The method according to claim 1 or 3, characterized in that Based on the mass of the hydrocracking catalyst containing graphite, the content of acidic cracking components is 10% to 40%, the content of alumina is 30% to 55%, the content of graphite is 2% to 10%, and the content of binder is less than 10%; and / or, based on the mass of the hydrocracking catalyst containing graphite, the content of Group VIII metal calculated as oxide is 1% to 8%, and the content of Group VIB metal calculated as oxide is 15% to 30%.

5. The method according to claim 1, wherein The hydrogen replacement is completed when the hydrogen gas volume fraction in the circulating hydrogen is adjusted to 95% to 99%.

6. The method according to claim 5, characterized in that When new hydrogen is introduced into the circulating hydrogen during the hydrogen replacement process, the inlet temperature of the second hydrocracking reaction zone is reduced to the same as the inlet temperature of the hydrofining reaction zone.

7. The method according to claim 1, characterized in that In the first sulfurization stage, the inlet temperatures of the hydrofining reaction zone and the second hydrocracking reaction zone are increased to a first constant temperature at a heating rate of 5-10°C / h.

8. The method according to claim 1, characterized in that In the second sulfiding stage, the inlet temperatures of the hydrofining reaction zone and the second hydrocracking reaction zone are increased to a second constant temperature at a heating rate of 3-10°C / h.

9. The method according to claim 1, characterized in that In the second vulcanization stage, the second constant temperature is 340-370°C.

10. The method according to claim 1, characterized in that The wet sulfiding process allows hydrogen sulfide to penetrate the catalyst bed before entering the first sulfiding stage. The process is as follows: low-nitrogen oil is introduced from the inlet of the hydrofining reaction zone to wet the catalyst, a closed-loop circulation of the low-nitrogen oil is established, and the inlet temperatures of the hydrofining reaction zone and the second hydrocracking reaction zone are increased to 190-210°C at a rate of ≯15°C / h. Then, a sulfiding agent is gradually injected into the low-nitrogen oil until hydrogen sulfide penetrates the entire catalyst bed. The nitrogen content of the low-nitrogen oil is ≯300 ppm by mass.

11. The method according to claim 10, characterized in that The conditions for hydrogen sulfide to penetrate the entire catalyst bed are as follows: the mass concentration of hydrogen sulfide in the circulating hydrogen is above 1000 ppm.

12. The method according to claim 10, characterized in that The conditions for hydrogen sulfide to penetrate the entire catalyst bed are as follows: the mass concentration of hydrogen sulfide in the circulating hydrogen is 1000~5000ppm.

13. The method according to claim 1, wherein After the first constant temperature sulfiding process is completed, the passivating agent is injected for passivation, and the injection rate of the sulfiding agent is adjusted so that the mass concentration of hydrogen sulfide in the circulating hydrogen reaches 5000-10000 ppm. The inlet temperature of the two-stage reactor is steadily increased at a rate of ≯8°C / h. Before ammonia breakthrough, the maximum temperature of the catalyst bed in the first hydrocracking reaction zone and the second hydrocracking reaction zone is controlled to be ≯260°C. When the mass content of ammonia in the high-density acidic water is greater than 1000 ppm, it is considered that ammonia has penetrated the catalytic bed. After ammonia breakthrough, the passivating agent injection rate is adjusted to maintain the mass concentration of ammonia in the high-density acidic water ≮5000 ppm, and the addition of the passivating agent is stopped until the constant temperature sulfiding process is completed. The passivating agent is selected from one or more of anhydrous liquid ammonia, cyclohexylamine, aniline, and n-butylamine.

14. The method according to claim 13, characterized in that After ammonia penetrates, the passivator injection rate is adjusted to maintain the ammonia concentration in the highly acidic water at 5000~15000 ppm until the constant temperature vulcanization process is completed and the passivator addition is stopped.

15. The method according to claim 13, characterized in that The condition for ammonia breakthrough is that the mass concentration of ammonium ions in the acidic water discharged from the bottom of the high-pressure separator reaches above 1000 ppm.

16. The method according to claim 15, characterized in that The condition for ammonia breakthrough is that the mass concentration of ammonium ions in the acidic water discharged from the bottom of the high-pressure separator is 1000-20000 ppm.

17. The method according to claim 10, wherein: The vulcanizing agent is at least one of dimethyl disulfide and carbon disulfide.

18. The method according to claim 10, wherein: The low-nitrogen oil is atmospheric distillate oil with an initial distillation point of 150-230° C., a final distillation point ≯360° C., and a water content ≯100 ppm.

19. The method according to claim 1, wherein The raw oil is selected from wax oil and / or diesel raw materials, the initial distillation point of the raw material is 155~380℃, the final distillation point is 390~750℃, the sulfur content is 0.2wt%~3.0wt%, and the nitrogen mass content is 500ppm~2000ppm.

20. The method according to claim 1, wherein The two-stage hydrocracking method is used to produce heavy naphtha, and the process is as follows: The feedstock oil is mixed with hydrogen and sequentially enters the first hydrorefining reaction zone and the first hydrocracking reaction zone. The resulting reaction effluent is passed through a fractionation system to obtain gas, light naphtha, heavy naphtha products, and a tail oil fraction. The tail oil fraction and hydrogen enter the second hydrocracking reaction zone of the second stage for reaction. The resulting effluent from the second hydrocracking reaction zone enters a fractionation system to obtain gas, light naphtha, heavy naphtha products, and a tail oil fraction. The initial distillation point of the tail oil fraction is 160-370°C.

21. The method according to claim 20, characterized in that The effluent from the second hydrocracking reaction zone and the effluent from the first hydrocracking reaction zone share a fractionation system.

22. The method according to claim 20, characterized in that During normal production, the reaction conditions in the first stage hydrofining reaction zone and the first cracking reaction zone are as follows: reaction pressure of 6.0-20.0 MPa; average reaction temperature of 270-440°C; hydrogen-to-oil volume ratio of 400:1-2200:1; hydrofining volume space velocity of 0.2-5.0 h -1 ; Hydrocracking volume space velocity is 0.5~5.0h -1 ; And / or, the reaction conditions of the second hydrocracking reaction zone in the second stage are as follows: reaction pressure of 6.0-20.0 MPa; average reaction temperature of 230-400°C; total volume space velocity of 0.5-4.0 h -1 ; The volume ratio of hydrogen to oil is 400:1 to 2200:

1.

23. The method according to claim 22, characterized in that During normal production, the reaction conditions in the first stage hydrofining reaction zone and the first cracking reaction zone are as follows: reaction pressure 9.0-16.0 MPa; average reaction temperature 320-415°C; hydrogen-to-oil volume ratio 600:1-1400:1; hydrofining volume space velocity 0.5-4.0 h -1 ; Hydrocracking volume space velocity is 0.5~3.0 h -1 ; And / or, the reaction conditions of the second hydrocracking reaction zone in the second stage are as follows: reaction pressure of 9.0-16.0 MPa; average reaction temperature of 300-360°C; total volume space velocity of 1.0-4.0 h -1 ; The volume ratio of hydrogen to oil is 800:1~1600:1.

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