A method of starting work

Abstract

The present invention relates to the technical field of hydrocracking start-up, and discloses a start-up method. Wherein, the method includes: (1) loading an ammonia-passivated sulfurized hydrocracking catalyst into the reaction zone of a hydrocracking unit; (2) introducing hydrogen into the reaction zone of the hydrocracking unit, heating the catalyst bed temperature to 80-130°C, then introducing a diesel fraction and optionally vegetable oil and secondary processing oil to wet the catalyst bed, and establishing a closed loop; (3) heating the catalyst bed temperature in step (2) in sections, first heating the catalyst bed temperature to 200-240°C for insulation, then heating to 280-330°C and then switching the feedstock oil, continuing to heat up to the hydrocracking reaction temperature, and producing. The method has the characteristics of fast, safe and environmentally friendly start-up process, and can effectively improve the yield of the catalyst target product.

Description

Technical Field

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

[0002] With the continuous development of new energy around the world, the proportion of oil, as a high-quality fossil resource, used as vehicle fuel has shown a downward trend. However, with the vigorous development of the global chemical industry and the continuous improvement of people's material living standards, more demands have been put forward for the production of chemical raw materials through oil resources.

[0003] With the increasing demand for ethylene at home and abroad, naphtha is an important raw material for the production of ethylene. At the same time, naphtha can be used as a raw material for fertilizer, catalytic reforming, and solvent oil production. Therefore, it is of great significance to increase naphtha production. Hydrocracking is an important method to achieve the chain breaking of large molecular hydrocarbons to form high-quality short-chain hydrocarbons with low heteroatom content. At the same time, the hydrocracking of wax oil to produce more naphtha is also a hot topic in domestic and foreign research.

[0004] The active metal components of the hydrocracking catalyst will have a higher catalytic activity only when they are converted into a sulfided state. Therefore, it is very important to sulfidize the catalyst to give full play to the performance of the catalyst. At present, the sulfidation of hydrogenation catalysts is divided into three forms: in-situ sulfidation, in-situ activation with external sulfur loading, and ex-situ sulfidation. Among them, both the in-situ sulfidation and the in-situ activation with external sulfur loading will produce a large amount of hydrogen sulfide during the start-up phase, and there is a risk of hydrogen sulfide leakage. The sulfidation process takes a long time. At the same time, there are many factors that affect the catalyst sulfidation process. The sulfidation process at the device site can easily cause incomplete sulfidation and affect the activity of the catalyst. The ex-situ sulfidation method perfectly solves the above problems. The catalyst sulfidation step is completed during the catalyst production period to ensure the stability of the catalyst performance. At the same time, the generation of hazardous gases is reduced and the sulfidation time is saved during the start-up phase.

[0005] In order to ensure safety during the start-up phase of the hydrocracking catalyst, the molecular sieve needs to be passivated during the heating process to avoid the danger of over-cracking and overheating during the start-up phase. At present, the method commonly used by factories is to inject ammonia during the sulfidation process, or use high-nitrogen start-up oil to passivate the molecular sieve. However, no matter which method is used, it will increase the complexity of the start-up process. At the same time, liquid ammonia or nitrogen-containing compounds are highly irritating substances, and diffusion or leakage will cause environmental or safety risks. Using the off-site ammonia passivation method, the molecular sieve passivation work is completed in the catalyst plant, and there is no need to supplement ammonia injection or passivation at the start-up site, which will significantly simplify the start-up process and ensure the safety of equipment and personnel.

[0006] Producing naphtha through hydrocracking technology is an important means to increase naphtha production. At present, by adjusting the composition of the catalyst, changing the type of molecular sieve in the catalyst, optimizing the metal components of the catalyst, adjusting the grading of different catalysts and changing the process conditions of the device operation, the yield of naphtha can be improved to a certain extent and the added value of the product can be increased. However, complex operations are required during the catalyst preparation process or application process, and the production cost of the catalyst will be increased to a certain extent.

[0007] Patent application CN103059913A discloses a start-up pre-sulfurization method for hydrogenation catalysts. Although the consumption of sulfiding agent can be reduced by using the purified sulfur-containing exhaust gas from refineries for catalyst sulfidation, the composition of the sulfur-containing exhaust gas is complex. Although purification can remove high-content C1-C4 light hydrocarbons, a small amount of impurity components such as CO in it may cause poisoning of active metals on the catalyst, destroy the active phase structure of the catalyst, and affect the activity of the catalyst.

[0008] Patent application CN104593051A discloses a method for starting up a sulfidation-type hydrogenation catalyst, which uses an oil film generated by added organic matter to passivate the heat release of the sulfidation-type catalyst. Although this method can effectively solve the problem of concentrated heat release during the sulfidation process, it cannot passivate the molecular sieve for catalysts containing molecular sieves, and there is a risk of temperature runaway during startup.

[0009] Patent application CN103566963A discloses a method of introducing alkaline nitrides onto a catalyst at a low temperature stage, followed by in-situ sulfidation and activation. Although the cracking reaction can be controlled to a certain extent, the introduction of nitrides into the sulfided catalyst in the form of an aqueous solution will greatly affect the hydrogenation activity of the catalyst, resulting in the destruction of the active center of the catalyst.

[0010] Patent application CN107446616A discloses a method of loading low molecular weight nitrides onto a conventional hydrocracking catalyst or introducing low molecular weight nitrides during the catalyst kneading and molding process to alleviate the risk of temperature runaway during the start-up of the hydrocracking unit. However, the catalyst also needs to be sulfurized during the start-up process, and the sulfurizing agent needs to be continuously injected during the start-up process to sulfurize the catalyst, which may result in hydrogen sulfide leakage, long start-up time, certain risks and high costs.

[0011] Patent application CN101210195A discloses a hydrocracking method for producing more chemical light oil from inferior heavy raw materials. The method uses the optimization of catalysts to process inferior heavy raw materials in a single-stage series of two reactors through a single-pass process, and the naphtha yield is about 35%. The process used in this method will limit the naphtha yield and it is difficult to further increase it.

[0012] Patent application CN104560169A discloses a hydrocracking method for producing heavy naphtha from high-nitrogen raw materials. The method is characterized in that the tail oil fraction produced after a primary reaction is pressurized and then re-enters the second reaction zone for a hydrocracking reaction, thereby improving the yield of the naphtha fraction to a certain extent. However, the zoning and enhanced conversion of hydrocarbons is not fully considered. In addition to the improvement of the target product, the yield of products with lower added value, such as dry gas and liquefied gas, will also be improved to a certain extent.

[0013] Patent application CN109777512A discloses a hydrocracking method for improving the yield of heavy naphtha, in which the circulating oil in the product is injected into different positions in the cracking reactor for continued cracking reaction. This method will significantly increase the complexity of the process, and puts forward requirements for the control of different bed temperatures and the control of the injection amount of hydrogen, resulting in the overall complexity of the process. Summary of the invention

[0014] The purpose of the present invention is to overcome the problems that the prior art requires online sulfidation and passivation steps during the hydrocracking transposition start-up process, which has certain safety and environmental risks, and the selectivity of the catalyst is difficult to improve, and the yield of the target product is low. A start-up method is provided, which has the characteristics of fast, safe and environmentally friendly start-up process, and can effectively improve the yield of the catalyst target product.

[0015] In order to achieve the above object, the present invention provides a start-up method, wherein the method comprises:

[0016] (1) loading an ammonia-passivated sulfided hydrocracking catalyst into a reaction zone of a hydrocracking unit;

[0017] (2) introducing hydrogen into the reaction zone of the hydrocracking unit to raise the temperature of the catalyst bed to 80-130° C., and then introducing diesel fraction and optionally vegetable oil and secondary process oil to wet the catalyst bed to establish a closed loop;

[0018] (3) The catalyst bed temperature of step (2) is raised in stages. The catalyst bed temperature is first raised to 200-240°C for insulation, then raised to 280-330°C, the feedstock oil is switched, and the temperature is further raised to the hydrocracking reaction temperature for production.

[0019] The inventors of the present invention have found in their research that by using the low-nitrogen diesel fraction of the present invention and optionally vegetable oil and secondary processed oil in combination as the start-up oil, and starting the process with the hydrocracking catalyst of the present invention as the catalyst, on the one hand, the process of catalyst sulfidation and ammonia passivation can be avoided, and the start-up steps can be simplified. After the catalyst is loaded, the oil can be directly added to increase the temperature for start-up. On the other hand, by optimizing and adjusting the start-up oil, a good match with the performance of the catalyst is achieved. By adjusting the process parameters of the start-up process, the active phase of the ammonia-passivated sulfided catalyst is modified through the high hydrogen deficiency characteristics of aromatic substances and unsaturated hydrocarbon substances in the start-up oil, the catalyst selectivity is improved, and the goal of improving the yield of the target product is achieved. DETAILED DESCRIPTION

[0020] The endpoints and any values ​​of the ranges disclosed in this article 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 each range, the endpoint values ​​of each range and the individual point values, and the individual point values ​​can be combined with each other to obtain one or more new numerical ranges, which should be considered as specifically disclosed in this article.

[0021] The present invention provides a method for starting a work, wherein the method comprises:

[0022] (1) loading an ammonia-passivated sulfided hydrocracking catalyst into a reaction zone of a hydrocracking unit;

[0023] (2) introducing hydrogen into the reaction zone of the hydrocracking unit, raising the temperature of the catalyst bed to 80-130° C., and then introducing diesel fraction and optionally vegetable oil to wet the catalyst bed to establish a closed loop;

[0024] (3) The catalyst bed temperature of step (2) is raised in stages. The catalyst bed temperature is first raised to 200-240°C for insulation, then raised to 280-330°C, the feedstock oil is switched, and the temperature is further raised to the hydrocracking reaction temperature for production.

[0025] By adopting the method of the present invention, on the one hand, the process of catalyst sulfidation and ammonia passivation can be avoided, and the start-up steps can be simplified. After the catalyst is loaded, oil can be directly added to heat up and start the operation. On the other hand, by optimizing and adjusting the start-up oil, good coordination with the performance of the catalyst is achieved. By adjusting the process parameters of the start-up process and taking advantage of the high hydrogen deficiency characteristics of aromatic hydrocarbon substances and unsaturated hydrocarbon substances in the start-up oil, the active phase of the ammonia passivated sulfided catalyst is modified, the catalyst selectivity is improved, and the target of improving the yield of the target product is achieved.

[0026] In the present invention, there is no particular limitation on the preparation method of the ammonia-passivated sulfurized hydrocracking catalyst in step (1). Preferably, in step (1), the ammonia-passivated sulfurized hydrocracking catalyst is prepared by the following method: the oxidized hydrocracking catalyst is first passivated outside the hydrocracking unit and then sulfurized. The method of the present invention is used to address the problems of catalyst sulfurization and acidic molecular sieve ammonia passivation during the start-up of existing hydrocracking catalysts. The ammonia passivation process is completed outside the unit to effectively avoid the potential dangers of poor passivation effect or ammonia injection during the start-up process, reduce the use and emission of toxic substances such as ammonia, and reduce the waste of passivators. It has the advantages of saving resources, low carbon and environmental protection, and avoids the characteristics of unstable passivation caused by high-nitrogen oil and the possible introduction of other impurities, and has an efficient and stable passivation effect. Furthermore, the method for completing the catalyst sulfurization process outside the unit provided by the present invention can effectively convert active metals into metal sulfides with high catalytic activity, has a good sulfurization effect, and avoids the use of sulfurization on site inside the unit. The safety and environmental risks brought by the use of sulfiding agents also avoid the risk of concentrated decomposition and over-temperature of sulfides that may occur during the activation of the sulfur-carrying catalyst in the device, while reducing the damage of the sulfiding agent to the human body and the investment in sulfur injection equipment; further, the method of the present invention adds an organic alcohol compound when preparing the ammonia passivation impregnation aqueous solution, which effectively increases the dispersibility of the nitrogen-containing compound on the catalyst, increases the retention, achieves a better passivation effect, and promotes the migration of the nitrogen-containing compound to the acidic position. At the same time, the addition of the organic alcohol compound is beneficial to the sulfurization process of the catalyst, effectively improving the activity of the catalyst. The oxidized catalyst is more conducive to the generation of an active phase with a high number of stacking layers during the sulfurization process, and has better hydrogenation activity.

[0027] In the present invention, the hydrocracking catalyst is a hydrocracking catalyst in a broad sense. In addition to conventional hydrocracking catalysts, the present invention may also include various molecular sieve-containing hydrocracking catalysts (e.g., hydro-upgrading catalysts) and specific hydrocracking catalysts known in the art. Preferably, in step (1), the hydrocracking catalyst comprises a cracking component, a hydrogenation component and a carrier.

[0028] In the present invention, there is no particular limitation on the type of cracking component, and it can be a cracking component conventionally defined in the art. Preferably, the cracking component includes an amorphous acidic component and / or a molecular sieve, the amorphous acidic component includes amorphous silicon aluminum and / or amorphous silicon magnesium, and the molecular sieve is at least one selected from Y-type molecular sieve, ZSM-5 molecular sieve, SAPO molecular sieve and MCM-41 mesoporous molecular sieve.

[0029] In the present invention, there is no particular limitation on the type of hydrogenation component, and it can be a cracking component conventionally defined in the art. Preferably, the hydrogenation component includes at least one of a Group VIII metal and a Group VIB metal, for example, the Group VIII metal is Co and / or Ni, and the Group VIB metal is Mo and / or W.

[0030] In the present invention, there is no particular limitation on the type of carrier, and the carrier may be any carrier conventionally defined in the art. Preferably, the carrier comprises a refractory porous material, and the refractory porous material is selected from at least one of aluminum oxide, silicon oxide, titanium oxide, magnesium oxide, zirconium oxide and activated carbon.

[0031] According to a preferred embodiment of the present invention, in step (1), based on the weight of the hydrocracking catalyst, the content of the cracking component is 10-60% by weight, the content of the carrier is 30-70% by weight, the content of the Group VIII metal calculated as oxide is 1-15% by weight, and the content of the Group VIB metal calculated as oxide is 5-30% by weight; further preferably, based on the weight of the hydrocracking catalyst, the content of the cracking component is 13-50% by weight, the content of the carrier is 35-65% by weight, the content of the Group VIII metal calculated as oxide is 3-12% by weight, and the content of the Group VIB metal calculated as oxide is 8-28% by weight.

[0032] According to a preferred embodiment of the present invention, in step (1), the passivation treatment comprises: in the presence of a solution containing an organic alcohol compound and a nitrogen-containing compound, an organic alcohol compound and an organic nitrogen-containing compound are loaded onto an oxidized hydrogenation catalyst by an impregnation method, and then a passivation treatment is performed, wherein the amount of the organic nitrogen-containing compound is such that the nitrogen content in the ammonia-passivated oxidized hydrocracking catalyst, calculated as an element, accounts for 0.1-8% of the weight of the oxidized hydrocracking catalyst before passivation, and the molar ratio of the organic alcohol compound to the metal atoms of Group VIII in the oxidized hydrocracking catalyst, calculated as a mole, is 0.1-3.

[0033] According to a preferred embodiment of the present invention, the amount of the organic nitrogen-containing compound and the solution containing the organic alcohol compound is such that the nitrogen content in the ammonia-passivated sulfided hydrocracking catalyst, calculated as an element, accounts for 0.5-5% of the weight of the sulfided hydrocracking catalyst before passivation, and the molar ratio of the organic alcohol compound to the metal atoms of Group VIII in the oxidized hydrocracking catalyst is 0.3-1.5.

[0034] According to a preferred embodiment of the present invention, the organic alcohol compound is selected from at least one of aliphatic alcohol compounds, alicyclic alcohol compounds and aromatic alcohol compounds. Preferably, the organic alcohol compound is selected from at least one of methanol, ethanol, ethylene glycol, propanol, propylene glycol, glycerol, butanediol, pentaerythritol and pentanepentol.

[0035] According to a preferred embodiment of the present invention, the organic alcohol compound has 1-15 carbon atoms, preferably 1-10 carbon atoms.

[0036] According to a preferred embodiment of the present invention, the organic nitrogen-containing compound is selected from at least one of alkylamine compounds, arylamine compounds, aniline compounds, methylaniline compounds, amide compounds, alcoholamine compounds and polyamine compounds, and is further preferably an alkylamine compound and / or an alcoholamine compound, for example, it can be selected from at least one of ethylenediamine, propylamine, butylamine, pentylamine, hexylamine, triethylamine, tert-butylamine, N,N-dihydroxyethylaniline, acetanilide, ethanolamine, diethanolamine, triethanolamine, diisopropanolamine, N-(2-hydroxyethyl)ethylenediamine, N-methyldiethanolamine, N,N-diisopropylethanolamine, 1,2-cyclohexanediamine, 1,3-propylenediamine, triethylenediamine, N,N-dimethyldipropylenetriamine, triethylenetetramine and hexamethylenetetramine.

[0037] According to a preferred embodiment of the present invention, the organic nitrogen-containing compound has 1-20 carbon atoms, preferably 2-15 carbon atoms.

[0038] In the present invention, a conventionally defined impregnation method may be selected, for example, saturated impregnation, unsaturated impregnation or supersaturated impregnation loading. Preferably, the loading may be carried out in at least one of the following ways, for example, the oxidized hydrogenation catalyst may be impregnated into a solution containing an organic alcohol compound and an organic nitrogen-containing compound; or the solution containing an organic alcohol compound and an organic nitrogen-containing compound may be sprayed onto the oxidized hydrogenation catalyst.

[0039] In the present invention, the passivation treatment includes drying the impregnated oxidized hydrogenation catalyst, and the drying conditions are not particularly limited in the present invention. Preferably, the drying conditions include: a temperature of 50-150°C and a time of 1-8 hours; more preferably, a temperature of 80-130°C and a drying time of 2-6 hours. In the present invention, there is no particular limitation on the pressure of the drying treatment, for example, it can be normal pressure.

[0040] In the present invention, preferably, the drying treatment is performed in an air atmosphere.

[0041] In the present invention, hydrogen is introduced into step (2) as a medium to increase the temperature of the catalyst bed. Preferably, in step (2), the heating rate of the catalyst bed is 5-25°C / h, more preferably 10-25°C / h.

[0042] In the present invention, there is no particular limitation on the temperature of the catalyst bed in step (2), as long as the start-up conditions are met, and those skilled in the art can make a selection according to actual needs. Preferably, in step (2), hydrogen is introduced to raise the temperature of the catalyst bed to 80-120°C. The advantage of adopting this preferred embodiment is that the catalyst has a certain ability to dissociate hydrogen in this temperature range, and hydrogen can be promoted to contact with a highly active catalyst in this process, and the generated adsorption heat can be released in advance, avoiding subsequent overheating and the released heat can be used for bed heating.

[0043] In the present invention, the type of diesel fraction in step (2) is limited to meet the start-up conditions. Preferably, in step (2), the diesel fraction is selected from at least one of straight-run diesel, vacuum light distillate oil and hydrocracked diesel, preferably straight-run diesel and / or hydrocracked diesel.

[0044] According to a preferred embodiment of the present invention, the distillation range of the diesel fraction is 210-380°C, preferably 210-370°C.

[0045] According to a preferred embodiment of the present invention, the nitrogen content in the diesel fraction is 100-800 ppm, preferably 100-500 ppm.

[0046] The selection of the diesel fraction according to the above preferred embodiment has the advantage of supplementing the passivation during the start-up process, and the clear distillation range is conducive to better miscibility with the organic matter loaded on the catalyst, so as to flush the catalyst bed.

[0047] In the present invention, diesel fractions, vegetable oils and secondary processed oils are selected to cooperate with each other as start-up oils to provide good conditions for hydrocracking processing. Preferably, in step (2), the vegetable oil is selected from at least one of corn oil, soybean oil, peanut oil, rapeseed oil, coconut oil, sunflower oil, olive oil and cottonseed oil, preferably at least one of corn oil, soybean oil, peanut oil, rapeseed oil, coconut oil, sunflower oil and cottonseed oil.

[0048] According to a preferred embodiment of the present invention, the distillation range of the vegetable oil is 230-580°C, preferably 230-560°C.

[0049] According to a preferred embodiment of the present invention, the bromine value of the vegetable oil is 8-20 gBr / 100 mL, preferably 10-15 gBr / 100 mL.

[0050] According to a preferred embodiment of the present invention, the secondary processing oil is selected from at least one of catalytic gasoline, pyrolysis gasoline, catalytic diesel and coker diesel.

[0051] According to a preferred embodiment of the present invention, the distillation range of the secondary processing oil is 170-380°C, preferably 190-370°C.

[0052] According to a preferred embodiment of the present invention, the density of the secondary processing oil is 0.85-0.98 g.cm -3 , preferably 0.89-0.96 g.cm -3 .

[0053] According to a preferred embodiment of the present invention, the aromatic content in the secondary processing oil is 50-90% by weight, preferably 60-90% by weight.

[0054] By selecting the high hydrogen deficiency characteristics of aromatic substances and unsaturated hydrocarbon substances in the vegetable oil and secondary processed oil under the above preferred embodiment, the active phase of the ammonia-passivated sulfided catalyst is modified to achieve the advantage of improving the catalyst selectivity and achieving the goal of increasing the yield of the target product.

[0055] According to a preferred embodiment of the present invention, in step (2), compared with 100 parts by weight of the hydrocracking catalyst, the amount of the diesel fraction is 200-2500 parts by weight, the amount of the vegetable oil is 0-250 parts by weight, and the amount of the secondary processing oil is 0-250 parts by weight; preferably, compared with 100 parts by weight of the hydrocracking catalyst, the amount of the diesel fraction is 300-2000 parts by weight, the amount of the vegetable oil is 20-200 parts by weight, and the amount of the secondary processing oil is 10-200 parts by weight. The advantage of adopting this preferred embodiment is that the ratio of polycyclic aromatic hydrocarbons and unsaturated hydrocarbons in the start-up oil is adjusted by matching the proportions of different oil products in the start-up oil, and the active phase can be better adjusted for different catalysts.

[0056] According to a preferred embodiment of the present invention, in step (2), the closed-loop circulation conditions include: pressure of 2-6 MPa, hydrogen-oil volume ratio of 500:1-1300:1, volume space velocity of 0.3-2.5 h -1 Preferably, the pressure is 3-6 MPa, the hydrogen-oil volume ratio is 600:1-1100:1, and the volume space velocity is 0.3-1.8 h -1 .

[0057] In the present invention, a staged heating method is used to heat the device. According to a preferred embodiment of the present invention, in step (3), the temperature of the catalyst bed in the previous stage is 210-230°C.

[0058] According to a preferred embodiment of the present invention, in step (3), the insulation time of the previous stage is 4-10 h, and the pressure is 2-6 MPa; preferably, the insulation time of the previous stage is 4-8 h, and the pressure is 3-6 MPa.

[0059] According to a preferred embodiment of the present invention, in step (3), the heating rate in the previous stage is 5-20°C / h, preferably 5-15°C / h.

[0060] The advantage of adopting this preferred embodiment is that within the low and medium pressure range, by gradually increasing the bed temperature and maintaining a constant temperature for a certain period of time, the active phase structure of the catalyst can be adjusted by utilizing the polycyclic aromatic hydrocarbons and unsaturated hydrocarbons in the feedstock oil under the condition of low hydrogen solubility, thereby modifying the carrier surface of the catalyst carrier.

[0061] According to a preferred embodiment of the present invention, in step (3), the pressure in the latter stage is 7-16 MPa, preferably 7-14 MPa.

[0062] According to a preferred embodiment of the present invention, in step (3), the catalyst bed temperature in the latter stage is 290-320°C.

[0063] According to a preferred embodiment of the present invention, in step (3), the heating rate of the catalyst bed in the latter stage is 5-15°C / h, preferably 5-12°C / h.

[0064] The advantage of adopting this preferred implementation mode is that the pressure increase enables the device to reach the subsequent normal reaction conditions, which is convenient for the subsequent switching of formal raw materials. At the same time, the temperature is raised under relatively slow conditions, which on the one hand avoids the rapid occurrence of hydrocracking reaction and overheating of the catalyst bed, and can also achieve adjustment and optimization of the overall performance of the catalyst under high temperature and high pressure.

[0065] In the present invention, there is no particular limitation on the type of feedstock oil, and it can be any commonly used processing feedstock oil in the art. Preferably, in step (3), the feedstock oil is selected from at least one of straight-run wax oil, straight-run diesel and catalytic diesel.

[0066] According to a preferred embodiment of the present invention, the crude oil can obtain a higher naphtha yield under the condition of hydrocracking, thereby maximizing the added value of the product.

[0067] In the present invention, the hydrocracking conditions can be selected in a wide range. In step (3), the density of the feedstock oil is 0.85-1 g.cm -3 , sulfur content 0.5-1.5wt%, nitrogen content 500-5000ppm, distillation range 300-500℃.

[0068] In the present invention, it is understood that the production conditions are the same or different from the conditions of the closed loop in step (2), preferably the same, and the present invention does not make specific restrictions. The production conditions described in the present invention include pressure, hydrogen-to-oil volume ratio and volumetric space velocity.

[0069] The specific embodiments of the present invention are described in detail below. It should be understood that the specific embodiments described herein are only used to illustrate and explain the present invention, and are not used to limit the present invention.

[0070] The properties of the start-up oil and feed oil used in the following examples and comparative examples are shown in Table 1.

[0071] Table 1

[0072]

[0073] To illustrate the features of the present invention, the examples and comparative examples are prepared from the same batch of industrially produced commercial oxidized hydrocracking catalysts having a Ni oxide content of 4.5 wt%, a W oxide content of 25.5 wt%, a Mo oxide content of 3.8 wt%, a Y-type molecular sieve content of 55 wt% and a balance of alumina. The oxidized catalysts are loaded with diethanolamine having a nitrogen content of 3.0 wt% and propylene glycol having an atomic ratio to Ni of 0.8. The catalysts are dried at 120° C., normal pressure and in flowing air for 4 hours to obtain an ammonia-passivated oxidized hydrocracking catalyst loaded with nitrogen and an organic alcohol. After dry sulfidation, the sulfidation is carried out using a combination of hydrogen sulfide and hydrogen gas, and the ammonia-passivated oxidized hydrocracking catalyst is sulfided outside the hydrocracking unit. The sulfidation conditions are: H2S volume fraction is 3%, H2 volume fraction is 97%, heating rate is 20°C / h, sulfidation temperature is 340°C, constant temperature time is 8h, sulfidation pressure is 4MPa, and gas-agent volume ratio is 400:1, to obtain an ammonia-passivated sulfided hydrocracking catalyst, and this catalyst is used in subsequent examples and comparative examples.

[0074] Example 1

[0075] The hydrocracking unit was filled with hydrocracking catalyst. After the unit was airtight, the circulating hydrogen compressor was started. The bed of the hydrocracking unit was heated by hydrogen at a heating rate of 25°C / h, so that the bed temperature was controlled at 120°C. Low-nitrogen start-up diesel A was introduced into the hydrocracking unit to wet and flush the bed. After the bed was fully wetted and flushed, the start-up diesel A, vegetable oil C and secondary processed oil E were switched to a closed-loop circulation, wherein the weight ratio of diesel A to catalyst was 20, the weight ratio of vegetable oil C to catalyst was 2, and the weight ratio of secondary processed oil E to catalyst was 2. The amount ratio is 2, the operating pressure is 6MPa, the bed layer is heated at a rate of 15℃ / h, and the temperature is kept constant when it reaches 230℃ for 8h. Then after the operating pressure is increased to 14MPa, the bed layer is heated at a rate of 12℃ / h. After the temperature reaches 320℃, the raw oil is switched. At the same time, the reaction temperature is gradually increased by adjusting the reaction conditions and the heating furnace until the product is qualified. At this time, the temperature is 355℃, and the reaction is officially started. The operating conditions of the closed-loop cycle and the reaction process are a pressure of 12MPa, a hydrogen-to-oil volume ratio of 700:1, and a volume space velocity of 0.8h -1 The product was subjected to simulated distillation. The specific test conditions were as follows: the sample was analyzed using the multidimensional gas chromatography SH / T 0558-2016 method for determining the boiling range distribution of petroleum fractions. The proportion of each fraction was determined according to the different distillation ranges. After the reaction was completed, the catalyst was characterized for unloading. The carbon content of the unloaded catalyst was characterized and analyzed. The specific test conditions were as follows: the carbon content in the catalyst was determined using HORIBA's EMIA-920V. The sample and flux were placed in a high-frequency induction furnace for oxygen combustion. The generated CO2 gas flowed through an infrared absorption cell to absorb infrared energy. The carbon content was obtained from the change in energy.

[0076] Example 2

[0077] The same hydrocracking catalyst as in Example 1 was filled into the hydrocracking unit. After the unit was airtight, the circulating hydrogen compressor was started, and the bed of the hydrocracking unit was heated by hydrogen at a heating rate of 10°C / h, so that the bed temperature was controlled at 80°C. Low-nitrogen start-up diesel A was introduced into the hydrocracking unit to wet and flush the bed. After the bed was fully wetted and flushed, the start-up diesel A, vegetable oil C and secondary processed oil E were switched to perform closed-loop circulation, wherein the weight ratio of diesel A to catalyst was 3, the weight ratio of vegetable oil C to catalyst was 0.20, and the weight ratio of secondary processed oil E to catalyst was 0.5. The weight ratio of the catalyst is 0.1, the operating pressure is 3MPa, the bed is heated at a rate of 5℃ / h, and the temperature is kept constant when it reaches 210℃ for 4h. Then the operating pressure is increased to 7MPa, the bed is heated at a rate of 5℃ / h, and the raw oil is switched after the temperature reaches 290℃. At the same time, the reaction temperature is gradually increased by adjusting the reaction conditions and the heating furnace until the product is qualified. At this time, the temperature is 358℃, and the reaction is officially started. The operating conditions of the closed-loop cycle and the reaction process are a pressure of 12MPa, a hydrogen-to-oil volume ratio of 700:1, and a volume space velocity of 0.8h -1 The product was subjected to simulated distillation according to the method of Example 1 to determine the proportion of each fraction. After the reaction was completed, the catalyst was characterized and the carbon content of the discharged catalyst was characterized and analyzed.

[0078] Example 3

[0079] The same hydrocracking catalyst as in Example 1 was filled into the hydrocracking unit. After the unit was airtight, the circulating hydrogen compressor was started, and the bed of the hydrocracking unit was heated by hydrogen at a heating rate of 15°C / h, so that the bed temperature was controlled at 100°C. Low-nitrogen start-up diesel A was introduced into the hydrocracking unit to wet and flush the bed. After the bed was fully wetted and flushed, the start-up diesel A, vegetable oil C and secondary processed oil E were switched to perform closed-loop circulation, wherein the weight ratio of the diesel A to the catalyst was 10, the weight ratio of the vegetable oil C to the catalyst was 1, and the weight ratio of the secondary processed oil E to the catalyst was 1. The weight ratio of the catalyst is 1, the operating pressure is 5MPa, the bed is heated at a rate of 10℃ / h, and the temperature is kept constant when it reaches 220℃ for 6h. Then after the operating pressure is increased to 10MPa, the bed is heated at a rate of 8℃ / h. After the temperature reaches 310℃, the raw oil is switched. At the same time, the reaction temperature is gradually increased by adjusting the reaction conditions and the heating furnace until the product is qualified. At this time, the temperature is 356℃, and the reaction is officially started. The operating conditions of the closed-loop cycle and the reaction process are a pressure of 12MPa, a hydrogen-to-oil volume ratio of 700:1, and a volume space velocity of 0.8h -1 The product was subjected to simulated distillation according to the method of Example 1 to determine the proportion of each fraction. After the reaction was completed, the catalyst was characterized and the carbon content of the discharged catalyst was characterized and analyzed.

[0080] Example 4

[0081] The same hydrocracking catalyst as in Example 1 was filled into the hydrocracking unit. After the unit was airtight, the circulating hydrogen compressor was started, and the bed of the hydrocracking unit was heated by hydrogen at a heating rate of 12°C / h, so that the bed temperature was controlled at 110°C. Low-nitrogen start-up diesel B was introduced into the hydrocracking unit to wet and flush the bed. After the bed was fully wetted and flushed, the start-up diesel B, vegetable oil C and secondary processed oil E were switched to perform closed-loop circulation, wherein the weight ratio of the diesel B used to the catalyst was 12, the weight ratio of the vegetable oil C to the catalyst was 1, and the weight ratio of the secondary processed oil E to the catalyst was 1. The weight ratio of the chemical agent is 1, the operating pressure is 4MPa, the bed is heated at a rate of 12℃ / h, and the temperature is kept constant when it reaches 220℃ for 6h. Then after the operating pressure is increased to 10MPa, the bed is heated at a rate of 10℃ / h. After the temperature reaches 310℃, the raw oil is switched. At the same time, the reaction temperature is gradually increased by adjusting the reaction conditions and the heating furnace until the product is qualified. At this time, the temperature is 360℃, and the reaction is officially started. The operating conditions of the closed-loop cycle and the reaction process are a pressure of 12MPa, a hydrogen-to-oil volume ratio of 700:1, and a volume space velocity of 0.8h -1 The product was subjected to simulated distillation according to the method of Example 1 to determine the proportion of each fraction. After the reaction was completed, the catalyst was characterized and the carbon content of the discharged catalyst was characterized and analyzed.

[0082] Example 5

[0083] The same hydrocracking catalyst as in Example 1 was filled into the hydrocracking unit. After the unit was airtight, the circulating hydrogen compressor was started, and the bed of the hydrocracking unit was heated by hydrogen at a heating rate of 18°C / h, so that the bed temperature was controlled at 100°C. Low-nitrogen start-up diesel A was introduced into the hydrocracking unit to wet and flush the bed. After the bed was fully wetted and flushed, the start-up diesel A, vegetable oil D and secondary processed oil E were switched to perform closed-loop circulation, wherein the weight ratio of the diesel A to the catalyst was 10, the weight ratio of the vegetable oil D to the catalyst was 1.5, and the weight ratio of the secondary processed oil E to the catalyst was 1.5. The weight ratio of the catalyst is 1.2, the operating pressure is 4MPa, the bed is heated at a rate of 8℃ / h, and the temperature is kept constant when it reaches 225℃ for 7h. Then the operating pressure is increased to 12MPa, the bed is heated at a rate of 10℃ / h, and the feedstock oil is switched after the temperature reaches 300℃. At the same time, the reaction temperature is gradually increased by adjusting the reaction conditions and the heating furnace until the product is qualified. At this time, the temperature is 362℃, and the reaction is officially started. The operating conditions of the closed-loop cycle and the reaction process are a pressure of 12MPa, a hydrogen-to-oil volume ratio of 700:1, and a volume space velocity of 0.8h -1 The product was subjected to simulated distillation according to the method of Example 1 to determine the proportion of each fraction. After the reaction was completed, the catalyst was characterized and the carbon content of the discharged catalyst was characterized and analyzed.

[0084] Example 6

[0085] The same hydrocracking catalyst as in Example 1 was filled into the hydrocracking unit. After the unit was airtight, the circulating hydrogen compressor was started, and the bed of the hydrocracking unit was heated by hydrogen at a heating rate of 15°C / h, so that the bed temperature was controlled at 110°C. Low-nitrogen start-up diesel A was introduced into the hydrocracking unit to wet and flush the bed. After the bed was fully wetted and flushed, the start-up diesel A, vegetable oil C and secondary processed oil F were switched to perform closed-loop circulation, wherein the weight ratio of the diesel A to the catalyst was 10, the weight ratio of the vegetable oil C to the catalyst was 1, and the weight ratio of the secondary processed oil F to the catalyst was 1. The weight ratio of the catalyst is 1, the operating pressure is 6MPa, the bed is heated at a rate of 10℃ / h, and the temperature is kept constant when it reaches 230℃ for 6h. Then after the operating pressure is increased to 12MPa, the bed is heated at a rate of 8℃ / h. After the temperature reaches 310℃, the raw oil is switched. At the same time, the reaction temperature is gradually increased by adjusting the reaction conditions and the heating furnace until the product is qualified. At this time, the temperature is 362℃, and the reaction is officially started. The operating conditions of the closed-loop cycle and the reaction process are a pressure of 12MPa, a hydrogen-to-oil volume ratio of 700:1, and a volume space velocity of 0.8h -1 The product was subjected to simulated distillation according to the method of Example 1 to determine the proportion of each fraction. After the reaction was completed, the catalyst was characterized and the carbon content of the discharged catalyst was characterized and analyzed.

[0086] Example 7

[0087] The same hydrocracking catalyst as in Example 1 was filled into the hydrocracking unit. After the unit was airtight, the circulating hydrogen compressor was started, and the bed of the hydrocracking unit was heated by hydrogen at a heating rate of 8°C / h, so that the bed temperature was controlled at 125°C. Low-nitrogen start-up diesel A was introduced into the hydrocracking unit to wet and flush the bed. After the bed was fully wetted and flushed, the start-up diesel A, vegetable oil C and secondary processed oil E were switched to perform closed-loop circulation, wherein the weight ratio of the diesel A to the catalyst was 2:2, the weight ratio of the vegetable oil C to the catalyst was 2.3, and the weight ratio of the secondary processed oil E to the catalyst was 2:1. The weight ratio of the catalyst is 2.3, the operating pressure is 2MPa, the bed is heated at a rate of 18℃ / h, and the temperature is kept constant when it reaches 235℃ for 9h. Then the operating pressure is increased to 10MPa, the bed is heated at a rate of 14℃ / h, and the raw oil is switched after the temperature reaches 280℃. At the same time, the reaction temperature is gradually increased by adjusting the reaction conditions and the heating furnace until the product is qualified. At this time, the temperature is 364℃, and the reaction is officially started. The operating conditions of the closed-loop cycle and the reaction process are a pressure of 12MPa, a hydrogen-to-oil volume ratio of 700:1, and a volume space velocity of 0.8h -1The product was subjected to simulated distillation according to the method of Example 1 to determine the proportion of each fraction. After the reaction was completed, the catalyst was characterized and the carbon content of the discharged catalyst was characterized and analyzed.

[0088] Example 8

[0089] The same hydrocracking catalyst as in Example 1 was filled into the hydrocracking unit. After the unit was gas-tight, the circulating hydrogen compressor was started. The bed of the hydrocracking unit was heated by hydrogen at a heating rate of 15°C / h, so that the bed temperature was controlled at 125°C. Diesel A was introduced into the hydrocracking unit to wet and flush the bed. After the bed was fully wetted and flushed, the bed was heated at a rate of 8°C / h at an operating pressure of 6MPa. When the temperature reached 210°C, the temperature was kept constant for 4h. Then, after the operating pressure was increased to 12MPa, the bed was heated at a rate of 10°C / h. After the temperature reached 310°C, the raw materials were switched. The reaction temperature was gradually increased by adjusting the reaction conditions and the heating furnace until the product was qualified. At this time, the temperature was 363°C, and the reaction was officially started. The operating conditions of the reaction process were a pressure of 12MPa, a hydrogen-to-oil volume ratio of 700:1, and a volume space velocity of 0.8h -1 The product was subjected to simulated distillation according to the method of Example 1 to determine the proportion of each fraction. After the reaction was completed, the catalyst was characterized and the carbon content of the discharged catalyst was characterized and analyzed.

[0090] Example 9

[0091] The same hydrocracking catalyst as in Example 1 was filled into the hydrocracking unit. After the unit was found to be airtight, the circulating hydrogen compressor was started. The bed of the hydrocracking unit was heated by hydrogen at a heating rate of 22°C / h, so that the bed temperature was controlled at 100°C. Low-nitrogen start-up diesel A was introduced into the hydrocracking unit to wet and flush the bed. After the bed was fully wetted and flushed, the start-up diesel A and the secondary processed oil E were switched to perform closed-loop circulation, wherein the weight ratio of the diesel A to the catalyst was 10, and the weight ratio of the secondary processed oil E to the catalyst was 1. The operating pressure is 5MPa, the bed layer is heated at a rate of 12℃ / h, and the temperature is kept constant when it reaches 220℃ for 8h. Then after the operating pressure is increased to 12MPa, the bed layer is heated at a rate of 10℃ / h. After the temperature reaches 300℃, the raw oil is switched. At the same time, the reaction temperature is gradually increased by adjusting the reaction conditions and the heating furnace until the product is qualified. At this time, the temperature is 364℃, and the reaction is officially started. The operating conditions of the closed-loop cycle and the reaction process are a pressure of 12MPa, a hydrogen-to-oil volume ratio of 700:1, and a volume space velocity of 0.8h -1 The product was subjected to simulated distillation according to the method of Example 1 to determine the proportion of each fraction. After the reaction was completed, the catalyst was characterized and the carbon content of the discharged catalyst was characterized and analyzed.

[0092] Comparative Example 1

[0093] The same hydrocracking catalyst as in Example 1 was filled into the hydrocracking unit. After the unit was gas-tight, the circulating hydrogen compressor was started, and the bed of the hydrocracking unit was heated by hydrogen at a heating rate of 15°C / h, so that the bed temperature was controlled at 100°C. The raw oil wax oil was introduced into the hydrocracking unit to wet and flush the bed. After the bed was fully wetted and flushed, the bed was heated at a rate of 12°C / h at an operating pressure of 5MPa. When the temperature reached 225°C, the temperature was kept constant for 6h. Then, after the operating pressure was increased to 10MPa, the bed was heated at a rate of 8°C / h. After the temperature reached 310°C, the reaction temperature was gradually increased by adjusting the reaction conditions and the heating furnace until the product was qualified. At this time, the temperature was 372°C, and the reaction was officially started. The operating conditions of the reaction process were a pressure of 12MPa, a hydrogen-to-oil volume ratio of 700:1, and a volume space velocity of 0.8h -1 The product was subjected to simulated distillation to determine the proportion of each fraction. After the reaction was completed, the catalyst was characterized and the carbon content of the unloaded catalyst was characterized and analyzed.

[0094] Comparative Example 2

[0095] The same hydrocracking catalyst as in Example 1 was filled into the hydrocracking unit. After the unit was gas-tight, the circulating hydrogen compressor was started, and the bed of the hydrocracking unit was heated by hydrogen at a heating rate of 8°C / h, so that the bed temperature was controlled at 120°C. Vegetable oil C was introduced into the hydrocracking unit to wet and flush the bed. After the bed was fully wetted and flushed, the bed was heated at a rate of 10°C / h at an operating pressure of 6MPa. When the temperature reached 220°C, the temperature was kept constant for 8h. Then, after the operating pressure was increased to 12MPa, the bed was heated at a rate of 10°C / h. After the temperature reached 315°C, the raw materials were switched. The reaction temperature was gradually increased by adjusting the reaction conditions and the heating furnace until the product was qualified. At this time, the temperature was 368°C, and the reaction was officially started. The operating conditions of the reaction process were a pressure of 12MPa, a hydrogen-to-oil volume ratio of 700:1, and a volume space velocity of 0.8h. -1 The product was subjected to simulated distillation to determine the proportion of each fraction. After the reaction was completed, the catalyst was characterized and the carbon content of the unloaded catalyst was characterized and analyzed.

[0096] Comparative Example 3

[0097] The same hydrocracking catalyst as in Example 1 was filled into the hydrocracking unit. After the unit was gas-tight, the circulating hydrogen compressor was started. The bed of the hydrocracking unit was heated by hydrogen at a heating rate of 20°C / h, so that the bed temperature was controlled at 110°C. Secondary processing oil E was introduced into the hydrocracking unit to wet and flush the bed. After the bed was fully wetted and flushed, the bed was heated at a rate of 8°C / h at an operating pressure of 5MPa. When the temperature reached 230°C, the temperature was kept constant for 6h. Then, after the operating pressure was increased to 13MPa, the bed was heated at a rate of 8°C / h. After the temperature reached 295°C, the raw materials were switched. The reaction temperature was gradually increased by adjusting the reaction conditions and the heating furnace until the product was qualified. At this time, the temperature was 368°C, and the reaction was officially started. The operating conditions of the reaction process were a pressure of 12MPa, a hydrogen-to-oil volume ratio of 700:1, and a volume space velocity of 0.8h -1 The product was subjected to simulated distillation to determine the proportion of each fraction. After the reaction was completed, the catalyst was characterized and the carbon content of the unloaded catalyst was characterized and analyzed.

[0098] Comparative Example 4

[0099] The same hydrocracking catalyst as in Example 1 was filled into the hydrocracking unit. After the unit was found to be airtight, the circulating hydrogen compressor was started, and the bed of the hydrocracking unit was heated by hydrogen at a heating rate of 20°C / h, so that the bed temperature was controlled at 115°C. Low-nitrogen start-up diesel A was introduced into the hydrocracking unit to wet and flush the bed. After the bed was fully wetted and flushed, the start-up vegetable oil C and the secondary processed oil E were switched to perform closed-loop circulation, wherein the weight ratio of the vegetable oil C to the catalyst was 2, and the weight ratio of the secondary processed oil E to the catalyst was 2. The operating pressure is 6MPa, the bed is heated at a rate of 10℃ / h, and the temperature is kept constant when it reaches 225℃ for 8h. Then, after the operating pressure is increased to 10MPa, the bed is heated at a rate of 12℃ / h. After the temperature reaches 290℃, the feed oil is switched. At the same time, the reaction temperature is gradually increased by adjusting the reaction conditions and the heating furnace until the product is qualified. At this time, the temperature is 366℃, and the reaction is officially started. The operating conditions of the closed-loop cycle and the reaction process are a pressure of 12MPa, a hydrogen-to-oil volume ratio of 700:1, and a volume space velocity of 0.8h -1 The product was subjected to simulated distillation to determine the proportion of each fraction. After the reaction was completed, the catalyst was characterized and the carbon content of the unloaded catalyst was characterized and analyzed.

[0100] Comparative Example 5

[0101] The same hydrocracking catalyst as in Example 1 was filled into the hydrocracking unit. After the unit was airtight, the circulating hydrogen compressor was started, and the bed of the hydrocracking unit was heated by hydrogen at a heating rate of 30°C / h, so that the bed temperature was controlled at 135°C. Low-nitrogen start-up diesel A was introduced into the hydrocracking unit to wet and flush the bed. After the bed was fully wetted and flushed, the start-up diesel A, vegetable oil C and secondary processed oil E were switched to perform closed-loop circulation, wherein the weight ratio of the diesel A to the catalyst was 28, the weight ratio of the vegetable oil C to the catalyst was 0.15, and the weight ratio of the secondary processed oil E to the catalyst was 2.5. The weight ratio of the catalyst is 0.15, the operating pressure is 8MPa, the bed is heated at a rate of 25℃ / h, and the temperature is kept constant when it reaches 260℃ for 12h. Then, after the operating pressure is increased to 12MPa, the bed is heated at a rate of 18℃ / h. After the temperature reaches 340℃, the raw oil is switched. At the same time, the reaction temperature is gradually increased by adjusting the reaction conditions and the heating furnace until the product is qualified. At this time, the temperature is 370℃, and the reaction is officially started. The operating conditions of the closed-loop cycle and the reaction process are a pressure of 12MPa, a hydrogen-to-oil volume ratio of 700:1, and a volume space velocity of 0.8h -1 The product was subjected to simulated distillation to determine the proportion of each fraction. After the reaction was completed, the catalyst was characterized and the carbon content of the unloaded catalyst was characterized and analyzed.

[0102] The effects of the above embodiments and comparative examples are listed in Table 2 for comparison, wherein, by analyzing the conditions that each fraction in the product reaches the corresponding product index (the initial distillation point -180°C fraction is naphtha, and the sulfur content is less than 0.5ppm and the nitrogen content is less than 0.5ppm), the reaction temperature of the device is examined. The lower the reaction temperature, the higher the activity of the catalyst; the higher the yield of naphtha, the better the selectivity of the catalyst; the carbon content of the catalyst represents the selectivity regulation effect of the catalyst and the deactivation of the catalyst by carbon deposition. If the carbon deposition on the catalyst is too low, it means that an effective active phase regulation effect cannot be formed, which will affect the catalyst selectivity to a certain extent; if the carbon deposition on the catalyst is too high, the surface of the active phase will be covered, resulting in too low activity of the catalyst.

[0103] Among them, the yield of naphtha (%) = weight of naphtha / total weight of product * 100%

[0104] Table 2

[0105]

[0106] It can be seen from the above embodiments and comparative examples that the start-up method overcomes the problems that the prior art requires online sulfurization and passivation steps during the hydrocracking transposition start-up process, which has certain safety and environmental risks, and the selectivity of the catalyst is difficult to improve, and the yield of the target product is low. The method has the characteristics of fast, safe and environmentally friendly start-up process, and can effectively improve the yield of the catalyst target product, improve the selectivity of the catalyst, and improve the naphtha yield of the catalyst.

[0107] The preferred embodiments of the present invention are described in detail above, but the present invention is not limited thereto. Within the technical concept of the present invention, the technical solution of the present invention can be subjected to a variety of simple modifications, including the combination of various technical features in any other suitable manner, and these simple modifications and combinations should also be regarded as the contents disclosed by the present invention and belong to the protection scope of the present invention.

Claims

1. A method for starting a work, characterized in that: The method includes: (1) Loading an ammonia-passivated sulfided hydrocracking catalyst into the reaction zone of a hydrocracking unit; (2) Hydrogen is introduced into the reaction zone of the hydrocracking unit to raise the temperature of the catalyst bed to 80-130°C, and then diesel fractions, vegetable oil and secondary processed oil are introduced to wet the catalyst bed to establish a closed loop; (3) The catalyst bed temperature of step (2) is raised in stages, firstly the catalyst bed temperature is raised to 200-240°C for insulation, then the temperature is raised to 280-330°C, the feedstock oil is switched, and the temperature is further raised to the hydrocracking reaction temperature for production; The nitrogen content in the diesel fraction is 100-800 ppm; The secondary processing oil is selected from at least one of catalytic gasoline, pyrolysis gasoline, catalytic diesel and coker diesel; In step (2), compared with 100 parts by weight of the hydrocracking catalyst, the amount of the diesel fraction is 300-2000 parts by weight, the amount of the vegetable oil is 20-200 parts by weight, and the amount of the secondary processing oil is 10-200 parts by weight; In step (3), the feedstock oil is selected from at least one of straight-run wax oil, straight-run diesel and catalytic diesel; The bromine value of the vegetable oil is 8-20 gBr / 100 mL.

2. The method according to claim 1, wherein: In step (1), the ammonia-passivated sulfided hydrocracking catalyst is prepared by the following method: the oxidized hydrocracking catalyst is first passivated and then sulfided outside the hydrocracking unit.

3. The method according to claim 2, wherein: The passivation treatment comprises: in the presence of a solution containing an organic alcohol compound and a nitrogen-containing compound, using an impregnation method to load the organic alcohol compound and the organic nitrogen-containing compound onto an oxidized hydrogenation catalyst, and then performing a passivation treatment; The amounts of the organic nitrogen-containing compound and the organic alcohol compound are such that the nitrogen content in the ammonia-passivated oxidized hydrocracking catalyst, calculated as an element, accounts for 0.1-8% of the weight of the oxidized hydrocracking catalyst before passivation, and the molar ratio of the organic alcohol compound to the metal atoms of Group VIII in the oxidized hydrocracking catalyst, calculated as a mole number, is 0.1-3.

4. The method according to claim 2 or 3, wherein: The vulcanization treatment is wet vulcanization or dry vulcanization.

5. The method according to claim 4, wherein: The sulfidation treatment comprises: in the presence of a sulfiding agent and hydrogen, the passivated catalyst is subjected to sulfidation treatment to obtain an ammonia-passivated sulfided hydrocracking catalyst.

6. The method according to claim 4, wherein: The vulcanizing agent in the wet vulcanization is polysulfide.

7. The method according to claim 4, wherein: The vulcanizing agent in the wet vulcanization is carbon disulfide and / or dimethyl disulfide.

8. The method according to claim 4, wherein: The vulcanizing agent in the dry vulcanization is selected from hydrogen sulfide and hydrogen.

9. The method according to claim 4, wherein: The dry vulcanization treatment conditions include: a heating rate of 5-100°C / h, a maximum vulcanization temperature of 240-400°C, a constant temperature time of 1-15h, a vulcanization pressure of 0.1-10MPa, and an aerosol volume ratio of 50-1000:

1.

10. The method according to claim 9, wherein: The conditions of the dry vulcanization treatment include: a heating rate of 10-80°C / h, a maximum vulcanization temperature of 260-380°C, a constant temperature time of 2-12h, a vulcanization pressure of 0.3-8MPa, and an aerosol volume ratio of 100-800:

1.

11. The method according to claim 4, wherein: The conditions of the wet vulcanization treatment include: a heating rate of 5-100°C / h, a maximum vulcanization temperature of 240-400°C, a constant temperature time of 1-15h, a vulcanization pressure of 0.1-10MPa, and a volume space velocity of 0.3-8h -1 , the volume ratio of hydrogen to oil is 100-1000:

1.

12. The method according to claim 11, wherein: The conditions of the wet vulcanization treatment include: a heating rate of 10-80°C / h, a maximum vulcanization temperature of 260-380°C, a constant temperature time of 2-12h, a vulcanization pressure of 0.3-8MPa, and a volume space velocity of 0.5-6h -1 , the volume ratio of hydrogen to oil is 150-800:

1.

13. The method according to any one of claims 1 to 3, wherein: In step (2), the heating rate of the catalyst bed is 5-25°C / h.

14. The method according to claim 13, wherein: In step (2), the heating rate of the catalyst bed is 10-25°C / h.

15. The method according to any one of claims 1 to 3, wherein: In step (2), hydrogen is introduced to raise the temperature of the catalyst bed to 80-120°C.

16. The method according to any one of claims 1 to 3, wherein: In step (2), the diesel fraction is selected from at least one of straight-run diesel, vacuum light distillate oil and hydrocracked diesel; And / or, the distillation range of the diesel fraction is 210-380°C; And / or, the nitrogen content in the diesel fraction is 100-500 ppm.

17. The method according to claim 16, wherein: In step (2), the diesel fraction is straight-run diesel and / or hydrocracked diesel.

18. The method according to claim 16, wherein: The distillation range of the diesel fraction is 210-370°C.

19. The method according to any one of claims 1 to 3, wherein: In step (2), the vegetable oil is selected from at least one of corn oil, soybean oil, peanut oil, rapeseed oil, coconut oil, sunflower oil, olive oil and cottonseed oil; And / or, the distillation range of the vegetable oil is 230-580°C; And / or, the distillation range of the secondary processing oil is 170-380°C; And / or, the density of the secondary processing oil is 0.85-0.98 g.cm -3 ; And / or, the aromatic content in the secondary process oil is 50-90 wt %.

20. The method according to claim 19, wherein: In step (2), the vegetable oil is selected from at least one of corn oil, soybean oil, peanut oil, rapeseed oil, coconut oil, sunflower seed oil and cottonseed oil.

21. The method according to claim 19, wherein: The distillation range of the vegetable oil is 230-560°C.

22. The method according to any one of claims 1 to 3, wherein: The bromine value of the vegetable oil is 10-15 gBr / 100 mL.

23. The method according to claim 19, wherein: The distillation range of the secondary processing oil is 190-370°C.

24. The method according to claim 19, wherein: The density of the secondary processing oil is 0.89-0.96 g.cm -3 .

25. The method of claim 19, wherein: The aromatic content in the secondary processing oil is 60-90% by weight.

26. The method according to any one of claims 1 to 3, wherein: In step (2), the closed-loop circulation conditions include: pressure of 2-6 MPa, hydrogen-oil volume ratio of 500:1-1300:1, volume space velocity of 0.3-2.5 h -1 .

27. The method according to claim 26, wherein: In step (2), the closed-loop circulation conditions include: pressure of 3-6 MPa, hydrogen-oil volume ratio of 600:1-1100:1, volume space velocity of 0.3-1.8 h -1 .

28. The method according to any one of claims 1 to 3, wherein: In step (3), the catalyst bed temperature in the first stage is 210-230°C; And / or, in step (3), the first stage insulation time is 4-10h and the pressure is 2-6MPa; And / or, in step (3), the heating rate in the first stage is 5-20°C / h.

29. The method according to claim 28, wherein: In step (3), the first stage insulation time is 4-8h and the pressure is 3-6MPa.

30. The method of claim 28, wherein: In step (3), the heating rate in the first stage is 5-15°C / h.

31. The method according to any one of claims 1 to 3, wherein: In step (3), the second stage pressure is 7-16 MPa; and / or, in step (3), the second stage catalyst bed temperature is 290-320°C; And / or, in step (3), the heating rate of the second stage catalyst bed is 5-15°C / h.

32. The method according to claim 31, wherein: In step (3), the second stage pressure is 7-14 MPa.

33. The method according to claim 31, wherein: In step (3), the heating rate of the second stage catalyst bed is 5-12°C / h.

34. The method according to any one of claims 1 to 3, wherein: The density of the raw oil is 0.85-1g.cm -3 , sulfur content 0.5-1.5wt%, nitrogen content 500-5000ppm, distillation range 300-500℃.

35. The method of claim 2, wherein: In step (1), the oxidized hydrocracking catalyst comprises a cracking component, a hydrogenation component and a carrier.

36. The method of claim 35, wherein: The cracking component includes an amorphous acidic component and / or a molecular sieve, the amorphous acidic component includes amorphous silicon aluminum and / or amorphous silicon magnesium, and the molecular sieve is at least one selected from Y-type molecular sieve, ZSM-5 molecular sieve, SAPO molecular sieve and MCM-41 mesoporous molecular sieve.

37. The method of claim 35, wherein: The hydrogenation component includes at least one of a Group VIII metal and a Group VIB metal.

38. The method of claim 35, wherein: The carrier comprises a refractory porous substance, and the refractory porous substance is selected from at least one of aluminum oxide, silicon oxide, titanium oxide, magnesium oxide, zirconium oxide and activated carbon.

39. The method of claim 37, wherein: In step (1), based on the weight of the oxidized hydrocracking catalyst, the content of the cracking component is 10-60% by weight, the content of the carrier is 30-70% by weight, the content of the Group VIII metal calculated as oxide is 1-15% by weight, and the content of the Group VIB metal calculated as oxide is 5-30% by weight.

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