CN117946739A

CN117946739A - Start-up method

Info

Publication number: CN117946739A
Application number: CN202211333410.9A
Authority: CN
Inventors: 翟维明; 刘锋; 晋超; 褚阳; 杨平; 杨清河
Original assignee: Sinopec Petrochemical Research Institute Co ltd; China Petroleum and Chemical Corp
Current assignee: Sinopec Petrochemical Research Institute Co ltd; China Petroleum and Chemical Corp
Priority date: 2022-10-28
Filing date: 2022-10-28
Publication date: 2024-04-30

Abstract

The invention relates to the technical field of hydrocracking startup, and discloses a startup method. Wherein the method comprises the following steps: (1) Filling ammonia passivation type sulfuration hydrocracking catalyst into a reaction zone of the hydrocracking device; (2) Introducing hydrogen into a reaction zone of the hydrocracking device, heating the catalyst bed to 80-130 ℃, then introducing diesel fraction and optionally vegetable oil and secondary processing oil to wet the catalyst bed, and establishing closed cycle; (3) And (3) raising the temperature of the catalyst bed in the step (2) in a sectional manner, firstly raising the temperature of the catalyst bed to 200-240 ℃ for heat preservation, then raising the temperature to 280-330 ℃, switching raw oil, and continuing to raise the temperature to the hydrocracking reaction temperature for production. The method has the characteristics of quick, safe and environment-friendly start-up process, and can effectively improve the yield of the target product of the catalyst.

Description

Start-up method

Technical Field

The invention relates to the technical field of hydrocracking startup, in particular to a startup method.

Background

With the continuous development of new global energy, the proportion of petroleum as a high-quality fossil resource used as vehicle fuel is in a continuous decreasing trend, but with the vigorous development of the global chemical industry and the continuous improvement of the living standard of people, more demands are put forth for producing chemical raw materials through petroleum resources.

With the increasing demand of ethylene at home and abroad, naphtha is an important raw material for producing ethylene, and can be used as a raw material for fertilizer, catalytic reforming and solvent oil production, so that the promotion of naphtha production has important significance. Hydrocracking is an important method for realizing the formation of short-chain hydrocarbon with high quality and less heteroatom content by breaking macromolecular hydrocarbon chains, and meanwhile, the hydrocracking of wax oil to produce naphtha in a large amount is a hot spot in domestic and foreign research.

The active metal component of the hydrocracking catalyst has higher catalytic activity only when being converted into a vulcanized state, so that the catalyst is subjected to vulcanization treatment, which is important for fully playing the performance of the catalyst. At present, the sulfuration of hydrogenation catalysts is divided into three forms of in-device sulfuration, in-device activation of out-device sulfur carrier and out-device sulfuration. The method of vulcanizing in the device and activating in the sulfur carrier outside the device can generate a large amount of hydrogen sulfide at the start-up stage, the risk of hydrogen sulfide leakage exists, the vulcanizing process needs to consume a long time, meanwhile, the factors influencing the vulcanizing process of the catalyst are relatively large, the vulcanizing process on the site of the device is easy to cause incomplete vulcanization, and the activity of the catalyst is influenced. The method of outside vulcanization perfectly solves the problems, completes the vulcanization step of the catalyst during the production of the catalyst, ensures the stability of the catalyst performance, reduces the generation of dangerous gas at the start-up stage and saves the vulcanization time.

In order to ensure safety in the startup phase, the hydrocracking catalyst needs to passivate the molecular sieve in the heating process, so that the risk of overtemperature caused by excessive cracking in the startup phase is avoided, the mode frequently used in factories at present is ammonia injection in the vulcanization process or passivation of the molecular sieve by using high-nitrogen startup oil, but the complexity of the startup process is increased in either mode, meanwhile, liquid ammonia or nitrogen-containing compounds are substances with stronger irritation, and environmental protection or safety risks are caused by diffusion or leakage. The passivation of the molecular sieve is completed in the catalyst factory by using the external ammonia passivation mode, and the additional ammonia injection or passivation is not needed at the start-up site, so that the start-up process can be obviously simplified, and the safety of equipment and personnel is ensured.

The production of naphtha by hydrocracking technology is an important means for increasing the yield of naphtha, and the yield of naphtha can be improved to a certain extent 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 modes of different catalysts and changing the process conditions of the device operation, so that the added value of the product is increased. However, complicated operations are required in the catalyst preparation process or the application process, and the production cost of the catalyst is increased to some extent.

Patent application CN103059913a discloses a start-up presulfiding process for hydrogenation catalysts. The sulfur-containing exhaust gas of the refinery after purification is used for vulcanizing the catalyst, so that the consumption of a vulcanizing agent can be reduced, but the composition of the sulfur-containing exhaust gas is complex, and although the light hydrocarbon with higher content of C ₁-C₄ can be removed by purification, a small amount of impurity components such as CO and the like can cause poisoning of active metals on the catalyst, damage the active phase structure of the catalyst and influence the activity of the catalyst.

Patent application CN104593051a discloses a startup method of a sulfided hydrogenation catalyst, which uses an oil film generated by added organic matters to passivate the heat release of the sulfided catalyst, and can effectively solve the problem of concentrated heat release in the sulfiding process, but the catalyst containing molecular sieve cannot play a role in passivating the molecular sieve, and the startup risk is presented.

Patent application CN103566963a discloses a method of introducing basic nitrides onto a catalyst at a low temperature stage, then performing in-situ sulfiding and activation processes, while allowing a degree of control over the cracking reaction, introducing nitrides into sulfided catalysts in the form of aqueous solutions will have a significant impact on the hydrogenation activity of the catalyst, resulting in destruction of the active centers of the catalyst.

Patent application CN107446616a discloses a method for loading low molecular nitride on a conventional hydrocracking catalyst or introducing low molecular nitride in a catalyst kneading forming process, so as to relieve the risk of temperature runaway of a hydrocracking device in a startup process, but the catalyst also needs to be vulcanized in the startup process, and continuously injecting a vulcanizing agent to vulcanize the catalyst in the startup process, so that the problems of hydrogen sulfide leakage, long startup time, certain danger, high cost and the like exist.

Patent application CN101210195A discloses a hydrocracking method for producing more chemical light oil from inferior heavy raw material, which adopts the optimization of catalyst, and processes the inferior heavy raw material under the single-stage series connection of two reactors in one pass flow, and the naphtha yield is about 35%. The process flow adopted by the method can lead to limited naphtha yield and difficult further improvement.

Patent application CN104560169a discloses a hydrocracking method for producing heavy naphtha from high-nitrogen raw material, which is characterized in that the tail oil fraction produced by one-time reaction is pressurized and then enters the second reaction zone again to carry out hydrocracking reaction, so that the yield of naphtha fraction can be improved to a certain extent, but the partition strengthening conversion of hydrocarbons cannot be fully considered, besides the improvement of target products, the yield of products with lower added values such as dry gas, liquefied gas and the like can be improved to a certain extent.

Patent application CN109777512a discloses a hydrocracking method for improving the yield of heavy naphtha, in which circulating oil in the product is injected into different positions in a cracking reactor to carry out continuous cracking reaction, this way can significantly increase the complexity of the process, control the temperature of different beds, and the injection amount of hydrogen puts requirements on control, resulting in the overall complexity of the process.

Disclosure of Invention

The invention aims to overcome the problems that in the prior art, on-line sulfuration and passivation steps are needed in the hydrocracking transposed startup process, certain safety and environmental protection risks exist, the selectivity of a catalyst is difficult to improve, and the yield of a target product is low.

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

(1) Filling ammonia passivation type sulfuration hydrocracking catalyst into a reaction zone of the hydrocracking device;

(2) Introducing hydrogen into a reaction zone of the hydrocracking device, heating the catalyst bed to 80-130 ℃, then introducing diesel fraction and optionally vegetable oil and secondary processing oil to wet the catalyst bed, and establishing closed cycle;

(3) And (3) raising the temperature of the catalyst bed in the step (2) in a sectional manner, firstly raising the temperature of the catalyst bed to 200-240 ℃ for heat preservation, then raising the temperature to 280-330 ℃, switching raw oil, and continuing to raise the temperature to the hydrocracking reaction temperature for production.

The inventor of the invention discovers in the research that the low-nitrogen diesel oil fraction, optionally vegetable oil and secondary processing oil are mutually matched to be used as starting oil, and simultaneously the hydrocracking catalyst is used as a catalyst for starting, so that on one hand, the processes of catalyst vulcanization and ammonia passivation can be avoided, the starting step is simplified, the catalyst can be directly fed into oil for heating up for starting after filling, on the other hand, the good matching is realized with the performance of the catalyst through the optimization and adjustment of the starting oil, and on the other hand, the high-hydrogen deficiency characteristic of aromatic hydrocarbon substances and unsaturated hydrocarbon substances in the starting oil is utilized for modifying the active phase of ammonia passivation type vulcanization state catalysis, so that the selectivity of the catalyst is improved, and the aim of improving the yield of a target product is fulfilled.

Detailed Description

The endpoints and any values of the ranges disclosed herein are not limited to the precise range or value, and are understood to encompass values approaching those ranges or values. For numerical ranges, one or more new numerical ranges may be found between the endpoints of each range, between the endpoint of each range and the individual point value, and between the individual point value, in combination with each other, and are to be considered as specifically disclosed herein.

The invention provides a start-up method, which comprises the following steps:

(2) Introducing hydrogen into a reaction zone of the hydrocracking device, heating the catalyst bed to 80-130 ℃, then introducing diesel fraction and optionally vegetable oil to wet the catalyst bed, and establishing closed cycle;

The method of the invention can avoid the processes of catalyst vulcanization and ammonia passivation, simplify the start-up step, directly feed oil and heat up to start up after the catalyst filling is completed, realize good cooperation with the performance of the catalyst through the optimization and adjustment of the start-up oil, modify the active phase of ammonia passivation type vulcanization catalyst by utilizing the high hydrogen deficiency characteristic of aromatic hydrocarbon substances and unsaturated hydrocarbon substances in the start-up oil through the adjustment of the technological parameters of the start-up process, realize the improvement of the selectivity of the catalyst and fulfill the aim of improving the yield of target products.

In the present invention, the method for producing the ammonia-passivated sulfided hydrocracking catalyst in step (1) is not particularly limited. Preferably, in the step (1), the ammonia-passivated sulfided hydrocracking catalyst is prepared by passivating the oxidized hydrocracking catalyst outside the hydrocracking apparatus and then sulfiding the oxidized hydrocracking catalyst. The method of the invention aims at the problems of catalyst vulcanization and acid molecular sieve ammonia passivation in the starting process of the existing hydrocracking catalyst, effectively avoids the potential danger of poor passivation effect or ammonia injection in the starting process, reduces the use and emission of toxic substances such as ammonia gas and the like, reduces the waste of passivating agent, has the advantages of saving resources, being low-carbon and environment-friendly, avoiding the unstable passivation brought by high nitrogen oil and possibly introducing other impurities, and has the characteristics of high-efficiency and stable passivation effect; furthermore, the method for completing the vulcanization process of the catalyst outside the device can effectively convert active metals into metal sulfides with high catalytic activity, has good vulcanization effect, avoids the safety and environmental protection risks caused by using vulcanizing agents on the vulcanization site in the device, also avoids the occurrence of the risk of concentrated decomposition and overtemperature of sulfides possibly occurring in the activation in the sulfur-carrying catalyst, and simultaneously reduces the damage of the vulcanizing agents to human bodies and the investment of sulfur injection equipment; furthermore, the method of the invention adds the organic alcohol compound when preparing the ammonia passivation soaking water solution, effectively increases the dispersibility of the nitrogen compound on the catalyst, increases the retention degree, plays a better passivation effect, promotes the migration of the nitrogen compound to the acid position, and simultaneously, the addition of the organic alcohol compound is beneficial to the vulcanization process of the catalyst, effectively improves the activity of the catalyst, and the oxidation state catalyst is more beneficial to generating an active phase with a high stacking layer number in the vulcanization process, thereby having better hydrogenation activity.

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

In the present invention, the kind of the cracking component is not particularly limited, and may be a cracking component conventionally defined in the art. Preferably, the cracking component comprises an amorphous acidic component comprising amorphous silica alumina and/or amorphous silica magnesia and/or a molecular sieve selected from at least one of a Y-type molecular sieve, a ZSM-5 molecular sieve, a SAPO molecular sieve, and an MCM-41 mesoporous molecular sieve.

In the present invention, the kind of the hydrogenation component is not particularly limited, and may be a cracking component conventionally defined in the art. Preferably, the hydrogenation component comprises at least one of a group VIII metal and a group VIB metal, e.g., co and/or Ni for the group VIII metal and Mo and/or W for the group VIB metal.

In the present invention, the kind of the carrier is not particularly limited, and may be a carrier conventionally defined in the art. Preferably, the support comprises a refractory porous material selected from at least one of alumina, silica, titania, magnesia, zirconia, and activated carbon.

According to a preferred embodiment of the present invention, in step (1), the cracking component is present in an amount of 10 to 60 wt.%, the support is present in an amount of 30 to 70 wt.%, the group VIII metal is present in an amount of 1 to 15 wt.%, and the group VIB metal is present in an amount of 5 to 30 wt.%, based on the weight of the hydrocracking catalyst; further preferably, the cracking component is present in an amount of from 13 to 50 wt.%, the support is present in an amount of from 35 to 65 wt.%, the group VIII metal is present in an amount of from 3 to 12 wt.%, the group VIB metal is present in an amount of from 8 to 28 wt.%, based on the weight of the hydrocracking catalyst.

According to a preferred embodiment of the present invention, in step (1), the passivation treatment includes: in the presence of an organic alcohol compound and a nitrogen-containing compound solution, loading the organic alcohol compound and the organic nitrogen-containing compound onto an oxidation state hydrogenation catalyst by adopting an impregnation method, and then carrying out passivation treatment, wherein the organic nitrogen-containing compound is used in an amount such that the content of nitrogen in the ammonia passivation type oxidation state hydrocracking catalyst is 0.1-8% of the weight of the oxidation state hydrocracking catalyst before passivation, and the molar ratio of the organic alcohol compound to the VIII group metal atom in the oxidation state hydrocracking catalyst is 0.1-3.

According to a preferred embodiment of the present invention, the organic nitrogen-containing compound and the solution containing the organic alcohol compound are used in such an amount that the content of nitrogen in terms of elements in the ammonia-passivated sulfur-state hydrocracking catalyst is 0.5 to 5% by weight of the sulfur-state hydrocracking catalyst before passivation, and the molar ratio of the number of moles of the organic alcohol compound to the number of moles of the group VIII metal atoms in the oxidized-state hydrocracking catalyst is 0.3 to 1.5.

According to a preferred embodiment of the present invention, the organic alcohol compound is selected from at least one of an aliphatic alcohol compound, an alicyclic alcohol compound and an aromatic alcohol compound, and preferably, the organic alcohol compound is selected from at least one of methanol, ethanol, ethylene glycol, propanol, propylene glycol, glycerol, butanediol, pentaerythritol and pentanol.

According to a preferred embodiment of the present invention, the organic alcohol compound has a carbon number of 1 to 15, preferably 1 to 10.

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, alcohol amine compounds and polyamine compounds, more preferably alkylamine compounds and/or alcohol amine compounds, and may be selected from at least one of ethylenediamine, propylamine, butylamine, pentylamine, hexylamine, triethylamine, tert-butylamine, N-dihydroxyethylaniline, acetanilide, ethanolamine, diethanolamine, triethanolamine, diisopropanolamine, N- (2-hydroxyethyl) ethylenediamine, N-methyldiethanolamine, N-diisopropylethanolamine, 1, 2-cyclohexanediamine, 1, 3-propylenediamine, triethylenediamine, N-dimethyldipropylenetriamine, triethylenetetramine and hexamethylenetetramine, for example.

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

In the present invention, a conventionally defined impregnation method may be selected, and for example, saturated impregnation, unsaturated impregnation or supersaturated impregnation load may be used. Preferably, the supporting may be performed in at least one manner such as immersing the oxidation state hydrogenation catalyst in a solution containing an organic alcohol compound and an organic nitrogen-containing compound; the solution containing the organic alcohol compound and the organic nitrogen-containing compound may also be sprayed onto the oxidation state hydrogenation catalyst.

In the present invention, the passivation treatment includes drying the impregnated oxidation state hydrogenation catalyst, and the drying conditions are not particularly limited in the present invention. Preferably, the conditions of the drying process include: the temperature is 50-150 ℃ and the time is 1-8 hours; further preferably, the temperature is 80-130℃and the drying time is 2-6 hours. In the present invention, the pressure of the drying treatment is not particularly limited, and may be, for example, normal pressure.

In the present invention, the drying treatment is preferably performed under an air atmosphere.

In the invention, hydrogen is introduced into the step (2) as a medium to raise the temperature of the catalyst bed. Preferably, in step (2), the temperature rise rate of the catalyst bed is from 5 to 25℃per hour, more preferably from 10 to 25℃per hour.

In the present invention, the temperature of the catalyst bed in the step (2) is not particularly limited, and may be selected by those skilled in the art according to actual requirements as long as the start-up conditions are satisfied. Preferably, in the step (2), hydrogen is introduced to raise the temperature of the catalyst bed to 80-120 ℃. The catalyst has a certain capability of dissociating hydrogen in the temperature range, can promote the hydrogen to be contacted with the catalyst with high activity in the process, and releases the generated adsorption heat in advance, so that the subsequent overtemperature and the released heat can be used for bed heating.

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

According to a preferred embodiment of the invention, the diesel fraction has a distillation range of 210-380 ℃, preferably 210-370 ℃.

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

The diesel fraction with the preferred embodiment has the advantage of supplementing passivation in the starting process, and the definition of the distillation range is beneficial to better mutual solubility with organic matters loaded on the catalyst, so that the catalyst bed is washed.

In the invention, diesel oil fraction and vegetable oil are selected to be matched with secondary processing oil to be used as starting oil, thus providing 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, canola oil, coconut oil, sunflower oil, olive oil and cottonseed oil, preferably at least one of corn oil, soybean oil, peanut oil, canola oil, coconut oil, sunflower oil and cottonseed oil.

According to a preferred embodiment of the invention, the vegetable oil has a distillation range of 230-580 ℃, preferably 230-560 ℃.

According to a preferred embodiment of the invention, the vegetable oil has a bromine number of 8-20gBr/100mL, preferably 10-15gBr/100mL.

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.

According to a preferred embodiment of the invention, the secondary processing oil has a distillation range of 170-380 ℃, preferably 190-370 ℃.

According to a preferred embodiment of the invention, the secondary processing oil has a density of 0.85-0.98g.cm ^-3, preferably 0.89-0.96g.cm ^-3.

According to a preferred embodiment of the invention, the secondary processing oil has an aromatic content of 50 to 90 wt.%, preferably 60 to 90 wt.%.

The active phase of the ammonia passivation type sulfuration catalyst is modified by selecting the characteristic of high hydrogen deficiency of aromatic hydrocarbon substances and unsaturated hydrocarbon substances in the vegetable oil and the secondary processing oil under the preferred embodiment, so that the selectivity of the catalyst is improved, and the target advantage of improving the yield of target products is achieved.

According to a preferred embodiment of the present invention, in the step (2), the diesel fraction is used in an amount of 200 to 2500 parts by weight, the vegetable oil is used in an amount of 0 to 250 parts by weight, and the secondary processing oil is used in an amount of 0 to 250 parts by weight, relative to 100 parts by weight of the hydrocracking catalyst; preferably, the diesel fraction is used in an amount of 300 to 2000 parts by weight, the vegetable oil is used in an amount of 20 to 200 parts by weight, and the secondary processing oil is used in an amount of 10 to 200 parts by weight, compared to 100 parts by weight of the hydrocracking catalyst. The advantage of adopting this kind of preferred embodiment is that through the cooperation of different oil proportions in the operation oil, adjusted the proportion of polycyclic aromatic hydrocarbon and unsaturated hydrocarbon in the operation oil, can play better active phase regulation's effect to different catalysts.

According to a preferred embodiment of the present invention, in step (2), the closed cycle conditions include: the pressure is 2-6MPa, and the hydrogen oil volume ratio is 500:1-1300:1, the volume airspeed is 0.3-2.5h ^-1; preferably, the pressure is 3-6MPa, and the hydrogen-oil volume ratio is 600:1-1100:1, the volume space velocity is 0.3-1.8h ^-1.

In the invention, a sectional heating mode is selected to heat the device. According to a preferred embodiment of the invention, in step (3), the catalyst bed temperature of the previous stage is 210-230 ℃.

According to a preferred embodiment of the invention, in step (3), the previous stage has a holding time of 4-10 hours and a pressure of 2-6MPa; preferably, the heat preservation time of the previous stage is 4-8h, and the pressure is 3-6MPa.

According to a preferred embodiment of the invention, in step (3), the temperature increase rate in the previous stage is 5-20 ℃/h, preferably 5-15 ℃/h.

The advantage of adopting this kind of preferred embodiment is that in low pressure and medium pressure range, through the gradual increase of bed temperature and the constant temperature treatment of certain time, utilize polycyclic aromatic hydrocarbon and unsaturated hydrocarbon in the raw materials oil to adjust the active phase structure of catalyst under the lower state of hydrogen solubility, carry out the modification to the carrier surface of catalyst carrier.

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

According to a preferred embodiment of the invention, in step (3), the temperature of the catalyst bed in the latter stage is 290-320 ℃.

According to a preferred embodiment of the invention, in step (3), the temperature of the catalyst bed in the latter stage is raised at a rate of from 5 to 15 c/h, preferably from 5 to 12 c/h.

The advantage of adopting this kind of preferred embodiment is that pressure rise makes the device reach follow-up normal reaction condition, makes things convenient for the switching of follow-up formal raw materials, and the temperature is risen under comparatively slow condition simultaneously, avoids hydrocracking reaction's rapid emergence on the one hand, and the catalyst bed overtemperature still can realize adjusting the optimization to the overall performance of catalyst under high temperature and high pressure.

In the present invention, the type of the raw oil is not particularly limited, and any processing raw oil commonly used in the art may be used. Preferably, in step (3), the raw oil is at least one selected from straight-run wax oil, straight-run diesel oil and catalytic diesel oil.

According to a preferred embodiment of the invention, the raw oil can obtain more naphtha yield under the hydrocracking condition, so that the maximization of the added value of the product is realized.

In the present invention, the range of conditions for hydrocracking is wide. In the step (3), the density of the raw oil is 0.85-1g.cm ^-3, the sulfur content is 0.5-1.5wt%, the nitrogen content is 500-5000ppm, and the distillation range is 300-500 ℃.

In the present invention, it is understood that the conditions for production and the conditions for the closed cycle in step (2) are the same or different, preferably the same, and the present invention is not particularly limited. The conditions for the production described in the present invention include pressure, hydrogen-to-oil volume ratio and volume space velocity.

The following describes specific embodiments of the present invention in detail. It should be understood that the detailed description and specific examples, while indicating and illustrating the invention, are not intended to limit the invention.

The properties of the starting oil and the raw oil used in the following examples and comparative examples are shown in Table 1.

TABLE 1

To illustrate the features of the present invention, the commercial oxidation state hydrocracking catalyst comprising Ni oxide 4.5 wt%, W oxide 25.5 wt%, mo oxide 3.8 wt%, Y-type molecular sieve 55wt% and alumina for the rest was prepared through loading 3.0 wt% of diethanolamine with nitrogen content and glycerol with Ni atom ratio of 0.8 with the oxidation state catalyst, and drying the catalyst in flowing air at 120 deg.c and normal pressure for 4 hr to obtain ammonia deactivated oxidation state hydrocracking catalyst with nitrogen and organic alcohol. And then, carrying out dry vulcanization, namely vulcanizing the ammonia passivation type oxidation state hydrocracking catalyst outside the hydrocracking device by adopting a combined gas of hydrogen sulfide and hydrogen, wherein the vulcanization conditions are as follows: the volume fraction of H ₂ S is 3%, the volume fraction of H ₂ is 97%, the heating rate is 20 ℃/H, the vulcanization temperature is 340 ℃, the constant temperature time is 8H, the vulcanization pressure is 4MPa, and the gas-agent volume ratio is 400:1, obtaining ammonia passivation type sulfuration hydrocracking catalyst, and adopting the catalyst in the subsequent examples and comparative examples.

Example 1

Filling a hydrocracking catalyst into a hydrocracking device, starting a circulating hydrogen compressor after the device is airtight qualified, heating the bed of the hydrocracking device through hydrogen, wherein the heating rate is 25 ℃/h, controlling the temperature of the bed at 120 ℃, introducing low-nitrogen starting diesel A into the hydrocracking device to moisten and flush the bed, switching the starting diesel A, the vegetable oil C and the secondary processing oil E for closed cycle after the bed is fully moistened and flushed, wherein the weight ratio of the diesel A to the catalyst is 20, the weight ratio of the vegetable oil C to the catalyst is 2, the weight ratio of the secondary processing oil E to the catalyst is 2, the operating pressure is 6MPa, heating the bed at the rate of 15 ℃/h, keeping the temperature constant when the temperature reaches 230 ℃, keeping the temperature constant for 8h, then raising the operating pressure to 14MPa, heating the bed at the rate of 12 ℃/h, switching the raw oil after the temperature reaches 320 ℃, gradually raising the reaction temperature to be qualified by adjusting the reaction condition and the heating furnace, starting the reaction at the temperature to 355 ℃, formally carrying out the reaction, wherein the operating condition of the circulation and the reaction process is that the hydrogen volume ratio is 700 MPa: 1, a volume space velocity of 0.8h ^-1. The product is subjected to simulated distillation, and specific test conditions are as follows: analyzing a sample by adopting a measuring method of the boiling range distribution of the SH/T0558-2016 petroleum fraction by using a multidimensional gas chromatography, determining the proportion of each fraction according to different distillation ranges, carrying out catalyst unloading characterization on the catalyst after the reaction is finished, and carrying out characterization analysis on the carbon content of the unloaded catalyst, wherein the specific test conditions are as follows: the method comprises the steps of measuring the carbon content in a catalyst by using EMIA-920V of HORIBA company, putting a sample and a fluxing agent together into a high-frequency induction furnace, introducing oxygen for combustion, enabling generated CO ₂ gas to flow through an infrared absorption tank, absorbing infrared energy, and obtaining the carbon content by energy change.

Example 2

Filling the hydrocracking device with the same hydrocracking catalyst in the embodiment 1, after the device is airtight qualified, starting a circulating hydrogen compressor, heating the bed of the hydrocracking device through hydrogen at a heating rate of 10 ℃/h, controlling the temperature of the bed to 80 ℃, starting to introduce low-nitrogen starting diesel A into the hydrocracking device to moisten and rinse the bed, after the bed is fully moistened and rinsed, switching the starting diesel A, the vegetable oil C and the secondary processing oil E for closed cycle, wherein the weight ratio of the diesel A to the catalyst is 3, the weight ratio of the vegetable oil C to the catalyst is 0.20, the weight ratio of the secondary processing oil E to the catalyst is 0.1, the operating pressure is 3MPa, the temperature of the bed is raised at a rate of 5 ℃/h, the temperature is kept constant when reaching 210 ℃, the constant temperature is kept for 4h, after the operating pressure is raised to 7MPa, heating the bed at a rate of 5 ℃/h, switching the raw oil after the temperature reaches 290 ℃, and gradually increasing the reaction temperature to be qualified, starting to react at 358 ℃ through adjusting the reaction conditions and the heating furnace, and performing the closed cycle, wherein the operating pressure is 700 MPa, and the operating pressure is the closed cycle is 700 MPa: 1, a volume space velocity of 0.8h ^-1. The product was subjected to simulated distillation as in example 1, the proportions of the individual fractions were determined, the catalyst was subjected to catalyst removal characterization after the reaction was completed, and the carbon content of the removed catalyst was subjected to characterization analysis.

Example 3

Filling the hydrocracking device with the same hydrocracking catalyst in the embodiment 1, after the device is airtight qualified, starting a circulating hydrogen compressor, heating the bed of the hydrocracking device through hydrogen at a heating rate of 15 ℃/h, controlling the temperature of the bed to be 100 ℃, starting to introduce low-nitrogen starting diesel A into the hydrocracking device to moisten and rinse the bed, after the bed is fully moistened and rinsed, switching the starting diesel A, the vegetable oil C and the secondary processing oil E for closed cycle, wherein the weight ratio of the diesel A to the catalyst is 10, the weight ratio of the vegetable oil C to the catalyst is 1, the weight ratio of the secondary processing oil E to the catalyst is 1, the operating pressure is 5MPa, heating the bed at a speed of 10 ℃/h, keeping constant temperature when the temperature reaches 220 ℃, keeping the temperature for 6h, after the operating pressure is increased to 10MPa, heating the bed at a speed of 8 ℃/h, switching the raw oil after the temperature reaches 310 ℃, gradually increasing the reaction temperature to be qualified by adjusting the reaction condition and the heating furnace, starting to react at the temperature of 356 ℃, and performing the closed cycle, wherein the volume ratio of the secondary processing oil E to the operating pressure is 700 MPa, and the operating pressure is the closed cycle is 700 MPa: 1, a volume space velocity of 0.8h ^-1. The product was subjected to simulated distillation as in example 1, the proportions of the individual fractions were determined, the catalyst was subjected to catalyst removal characterization after the reaction was completed, and the carbon content of the removed catalyst was subjected to characterization analysis.

Example 4

Filling the hydrocracking device with the same hydrocracking catalyst in the embodiment 1, after the device is airtight qualified, starting a circulating hydrogen compressor, heating the bed of the hydrocracking device through hydrogen at a heating rate of 12 ℃/h, controlling the temperature of the bed to be 110 ℃, starting to introduce low-nitrogen starting diesel oil B into the hydrocracking device to moisten and rinse the bed, after the bed is fully moistened and rinsed, switching the starting diesel oil B, the vegetable oil C and the secondary processing oil E for closed cycle, wherein the weight ratio of the diesel oil B to the catalyst is 12, the weight ratio of the vegetable oil C to the catalyst is 1, the weight ratio of the secondary processing oil E to the catalyst is 1, the operating pressure is 4MPa, heating the bed at a speed of 12 ℃/h, keeping constant temperature when the temperature reaches 220 ℃, keeping constant temperature for 6h, after the operating pressure is increased to 10MPa, heating the bed at a speed of 10 ℃/h, switching the raw oil after the temperature reaches 310 ℃, gradually increasing the reaction temperature to be qualified by adjusting the reaction conditions and the heating furnace, starting to perform the reaction when the temperature is 360 ℃, the closed cycle is started, the volume ratio of the secondary processing oil E to be 700 MPa, and the operating pressure is 700 MPa: 1, a volume space velocity of 0.8h ^-1. The product was subjected to simulated distillation as in example 1, the proportions of the individual fractions were determined, the catalyst was subjected to catalyst removal characterization after the reaction was completed, and the carbon content of the removed catalyst was subjected to characterization analysis.

Example 5

Filling the hydrocracking device with the same hydrocracking catalyst in the embodiment 1, after the device is airtight qualified, starting a circulating hydrogen compressor, heating the bed of the hydrocracking device through hydrogen at a heating rate of 18 ℃/h, controlling the temperature of the bed to be 100 ℃, starting to introduce low-nitrogen starting diesel A into the hydrocracking device to moisten and rinse the bed, after the bed is fully moistened and rinsed, switching the starting diesel A, the vegetable oil D and the secondary processing oil E for closed cycle, wherein the weight ratio of the diesel A to the catalyst is 10, the weight ratio of the vegetable oil D to the catalyst is 1.5, the weight ratio of the secondary processing oil E to the catalyst is 1.2, the operating pressure is 4MPa, the temperature of the bed is raised at a speed of 8 ℃/h, the constant temperature is carried out when the temperature reaches 225 ℃, the constant temperature is maintained for 7h, after the operating pressure is raised to 12MPa, heating the bed at a speed of 10 ℃/h, switching the raw oil after the temperature reaches 300 ℃, and gradually increasing the reaction temperature to be qualified, starting to react at a temperature of 362 ℃ by adjusting the reaction condition and the heating furnace, and performing the closed cycle under the reaction condition that the operating pressure is 700 MPa, and the operating pressure is the closed cycle is the hydrogen volume ratio of the product is 700 MPa). 1, a volume space velocity of 0.8h ^-1. The product was subjected to simulated distillation as in example 1, the proportions of the individual fractions were determined, the catalyst was subjected to catalyst removal characterization after the reaction was completed, and the carbon content of the removed catalyst was subjected to characterization analysis.

Example 6

Filling the hydrocracking device with the same hydrocracking catalyst in the embodiment 1, after the device is airtight qualified, starting a circulating hydrogen compressor, heating the bed of the hydrocracking device through hydrogen at a heating rate of 15 ℃/h, controlling the temperature of the bed to be 110 ℃, starting to introduce low-nitrogen starting diesel A into the hydrocracking device to moisten and rinse the bed, after the bed is fully moistened and rinsed, switching the starting diesel A, the vegetable oil C and the secondary processing oil F to perform closed cycle, wherein the weight ratio of the diesel A to the catalyst is 10, the weight ratio of the vegetable oil C to the catalyst is 1, the weight ratio of the secondary processing oil F to the catalyst is 1, the operating pressure is 6MPa, heating the bed at a speed of 10 ℃/h, keeping constant temperature when the temperature reaches 230 ℃, keeping constant temperature for 6h, then heating the operating pressure to 12MPa, heating the bed at a speed of 8 ℃/h, switching the raw oil after the temperature reaches 310 ℃, gradually increasing the reaction temperature to be qualified by adjusting the reaction conditions and the heating furnace, starting to perform the reaction at the temperature of 362 ℃, and performing the closed cycle, wherein the operating pressure is 700 MPa, and the operating pressure is the closed cycle is 700 MPa: 1, a volume space velocity of 0.8h ^-1. The product was subjected to simulated distillation as in example 1, the proportions of the individual fractions were determined, the catalyst was subjected to catalyst removal characterization after the reaction was completed, and the carbon content of the removed catalyst was subjected to characterization analysis.

Example 7

Filling the hydrocracking device with the same hydrocracking catalyst in the embodiment 1, after the device is airtight qualified, starting a circulating hydrogen compressor, heating the bed of the hydrocracking device through hydrogen at a heating rate of 8 ℃/h, controlling the temperature of the bed to be 125 ℃, starting to introduce low-nitrogen starting diesel A into the hydrocracking device to moisten and rinse the bed, after the bed is fully moistened and rinsed, switching the starting diesel A, the vegetable oil C and the secondary processing oil E for closed cycle, wherein the weight ratio of the diesel A to the catalyst is 22, the weight ratio of the vegetable oil C to the catalyst is 2.3, the weight ratio of the secondary processing oil E to the catalyst is 2.3, the operating pressure is 2MPa, the temperature of the bed is raised at a speed of 18 ℃/h, the constant temperature is carried out when the temperature reaches 235 ℃, the constant temperature is maintained for 9h, after the operating pressure is raised to 10MPa, heating the bed at a speed of 14 ℃/h, switching the raw oil after the temperature reaches 280 ℃, and gradually increasing the reaction temperature to be qualified by adjusting the reaction conditions and the heating furnace, starting to react at a temperature of 364 ℃ and a normal volume ratio of the secondary processing oil is 700 MPa, and the operating pressure is the closed cycle is 700 MPa. 1, a volume space velocity of 0.8h ^-1. The product was subjected to simulated distillation as in example 1, the proportions of the individual fractions were determined, the catalyst was subjected to catalyst removal characterization after the reaction was completed, and the carbon content of the removed catalyst was subjected to characterization analysis.

Example 8

Filling the hydrocracking device with the same hydrocracking catalyst in the embodiment 1, after the device is qualified in airtight, starting a circulating hydrogen compressor, heating the bed of the hydrocracking device through hydrogen, controlling the temperature of the bed to be 125 ℃ at a heating rate of 15 ℃/h, starting to introduce the diesel A wetting and flushing bed into the hydrocracking device, after the bed is fully wetted and flushed, heating the bed at a constant temperature of 8 ℃/h at a constant temperature of 210 ℃ for 4 hours under an operating pressure of 6MPa, heating the bed at a speed of 10 ℃/h after the operating pressure is increased to 12MPa, switching raw materials after the temperature reaches 310 ℃, gradually increasing the reaction temperature to be qualified by adjusting the reaction condition and a heating furnace, starting to formally react at the temperature of 363 ℃, and controlling the operating condition of the reaction process to be 12MPa with a hydrogen-oil volume ratio of 700:1, a volume space velocity of 0.8h ^-1. The product was subjected to simulated distillation as in example 1, the proportions of the individual fractions were determined, the catalyst was subjected to catalyst removal characterization after the reaction was completed, and the carbon content of the removed catalyst was subjected to characterization analysis.

Example 9

Filling the hydrocracking device with the same hydrocracking catalyst in the embodiment 1, after the device is qualified in airtight, starting a circulating hydrogen compressor, heating the bed of the hydrocracking device through hydrogen, controlling the temperature of the bed to be 100 ℃ at a heating rate of 22 ℃/h, starting to introduce low-nitrogen starting diesel A into the hydrocracking device to moisten and flush the bed, after the bed is fully moistened and flushed, switching the starting diesel A and secondary processing oil E to carry out closed cycle, wherein the weight ratio of the diesel A to the catalyst is 10, the weight ratio of the secondary processing oil E to the catalyst is 1, the operating pressure is 5MPa, the bed is heated at a speed of 12 ℃/h, the temperature is kept constant when the temperature reaches 220 ℃, the constant temperature is 8h, after the operating pressure is raised to 12MPa, the bed is heated at a speed of 10 ℃/h, switching the raw oil after the temperature reaches 300 ℃, and gradually increasing the reaction temperature to be qualified by adjusting the reaction condition and the heating furnace, at the moment, starting the closed cycle and the reaction process operating condition is that the pressure is 12MPa, and the volume ratio of the hydrogen is 700:1, a volume space velocity of 0.8h ^-1. The product was subjected to simulated distillation as in example 1, the proportions of the individual fractions were determined, the catalyst was subjected to catalyst removal characterization after the reaction was completed, and the carbon content of the removed catalyst was subjected to characterization analysis.

Comparative example 1

Filling the hydrocracking device with the same hydrocracking catalyst in the embodiment 1, after the device is qualified in airtight, starting a circulating hydrogen compressor, heating the bed of the hydrocracking device through hydrogen, controlling the temperature of the bed to be 100 ℃ at a heating rate of 15 ℃/h, starting to introduce raw oil wax oil into the hydrocracking device to wet and rinse the bed, after the bed is fully wetted and rinsed, heating the bed at a constant temperature of 12 ℃/h at an operating pressure of 5MPa, keeping the temperature at 225 ℃ for 6 hours, then heating the bed at a speed of 8 ℃/h after the operating pressure is increased to 10MPa, gradually increasing the reaction temperature to be qualified by adjusting the reaction condition and a heating furnace, starting to formally react at a temperature of 372 ℃, and keeping the operating condition of the reaction process at a pressure of 12MPa at a hydrogen oil volume ratio of 700:1, a volume space velocity of 0.8h ^-1. And (3) carrying out simulated distillation on the product, determining the proportion of each fraction section, carrying out catalyst unloading characterization on the catalyst after the reaction is finished, and carrying out characterization analysis on the carbon content of the unloaded catalyst.

Comparative example 2

Filling the hydrocracking device with the same hydrocracking catalyst in the embodiment 1, after the device is qualified in airtight, starting a circulating hydrogen compressor, heating the bed of the hydrocracking device through hydrogen at a heating rate of 8 ℃/h, controlling the temperature of the bed at 120 ℃, starting to introduce vegetable oil C into the hydrocracking device for wetting and flushing the bed, after the bed is fully wetted and flushed, heating the bed at a constant temperature of 10 ℃/h at a temperature of 220 ℃ for 8h under an operating pressure of 6MPa, heating the bed at a temperature of 10 ℃/h after the operating pressure is increased to 12MPa, switching raw materials after the temperature reaches 315 ℃, gradually increasing the reaction temperature to be qualified by adjusting the reaction condition and a heating furnace, starting to perform a formal reaction at the temperature of 368 ℃, wherein the operating condition of the reaction process is the pressure of 12MPa, and the hydrogen-oil volume ratio is 700:1, a volume space velocity of 0.8h ^-1. And (3) carrying out simulated distillation on the product, determining the proportion of each fraction section, carrying out catalyst unloading characterization on the catalyst after the reaction is finished, and carrying out characterization analysis on the carbon content of the unloaded catalyst.

Comparative example 3

Filling the hydrocracking device with the same hydrocracking catalyst in the embodiment 1, after the device is qualified in airtight, starting a circulating hydrogen compressor, heating the bed of the hydrocracking device through hydrogen at a heating rate of 20 ℃/h, controlling the temperature of the bed at 110 ℃, starting to introduce secondary processing oil E into the hydrocracking device for wetting and flushing the bed, after the bed is fully wetted and flushed, heating the bed at a constant temperature of 8 ℃/h at a constant temperature of 230 ℃ for 6h under an operating pressure of 5MPa, heating the bed at a rate of 8 ℃/h after the operating pressure is increased to 13MPa, switching raw materials after the temperature reaches 295 ℃, gradually increasing the reaction temperature to be qualified by adjusting the reaction condition and a heating furnace, starting to formally react at the temperature of 368 ℃, wherein the operating condition of the reaction process is that the pressure is 12MPa, and the hydrogen-oil volume ratio is 700:1, a volume space velocity of 0.8h ^-1. And (3) carrying out simulated distillation on the product, determining the proportion of each fraction section, carrying out catalyst unloading characterization on the catalyst after the reaction is finished, and carrying out characterization analysis on the carbon content of the unloaded catalyst.

Comparative example 4

Filling the hydrocracking device with the same hydrocracking catalyst in the embodiment 1, after the device is qualified in airtight, starting a circulating hydrogen compressor, heating the bed of the hydrocracking device through hydrogen at a heating rate of 20 ℃/h, controlling the temperature of the bed to 115 ℃, starting to introduce low-nitrogen starting diesel oil A into the hydrocracking device to moisten and flush the bed, after the bed is fully moistened and flushed, switching the starting vegetable oil C and the secondary processing oil E for closed cycle, wherein the weight ratio of the vegetable oil C to the catalyst is 2, the weight ratio of the secondary processing oil E to the catalyst is 2, the operating pressure is 6MPa, the bed is heated at a speed of 10 ℃/h, the constant temperature is carried out when the temperature reaches 225 ℃, the constant temperature time is 8h, after the operating pressure is increased to 10MPa, the bed is heated at a speed of 12 ℃/h, switching the raw oil after the temperature reaches 290 ℃, and gradually increasing the reaction temperature to be qualified by adjusting the reaction condition and the heating furnace, at the moment, starting the closed cycle and the reaction process operating condition is 12MPa, and the volume ratio of the hydrogen is 700 formally: 1, a volume space velocity of 0.8h ^-1. And (3) carrying out simulated distillation on the product, determining the proportion of each fraction section, carrying out catalyst unloading characterization on the catalyst after the reaction is finished, and carrying out characterization analysis on the carbon content of the unloaded catalyst.

Comparative example 5

Filling the hydrocracking device with the same hydrocracking catalyst in the embodiment 1, after the device is airtight qualified, starting a circulating hydrogen compressor, heating the bed of the hydrocracking device through hydrogen at a heating rate of 30 ℃/h, controlling the temperature of the bed to be 135 ℃, starting to introduce low-nitrogen starting diesel A into the hydrocracking device to moisten and rinse the bed, after the bed is fully moistened and rinsed, switching the starting diesel A, the vegetable oil C and the secondary processing oil E for closed cycle, wherein the weight ratio of the diesel A to the catalyst is 28, the weight ratio of the vegetable oil C to the catalyst is 0.15, the weight ratio of the secondary processing oil E to the catalyst is 0.15, the operating pressure is 8MPa, the temperature of the bed is raised at a speed of 25 ℃/h, the constant temperature is carried out when the temperature reaches 260 ℃, the constant temperature is maintained for 12h, after the operating pressure is raised to 12MPa, heating the bed at a speed of 18 ℃/h, switching the raw oil after the temperature reaches 340 ℃, and gradually increasing the reaction temperature to be qualified, starting to react at 370 ℃ through adjusting the reaction conditions and the heating furnace, and performing the closed cycle, wherein the operating pressure is 700 MPa, and the operating pressure is the closed cycle is 700 MPa. 1, a volume space velocity of 0.8h ^-1. And (3) carrying out simulated distillation on the product, determining the proportion of each fraction section, carrying out catalyst unloading characterization on the catalyst after the reaction is finished, and carrying out characterization analysis on the carbon content of the unloaded catalyst.

The effects of the above examples and comparative examples are listed in table 2 for comparison, wherein the reaction temperature of the apparatus is examined under the condition that the respective fractions in the product reach the qualification of the corresponding product index (the initial distillation point-180 ℃ C. Fraction is divided into naphtha, the sulfur content is less than 0.5ppm, and the nitrogen content is less than 0.5 ppm), and the lower the reaction temperature, the higher the activity of the catalyst is considered; the higher the naphtha yield, the better the selectivity of the catalyst; the carbon content of the catalyst represents the selective regulation effect and deactivation condition of carbon deposition of the catalyst, and the carbon deposition on the catalyst is too low, which indicates that the catalyst selectivity is affected to a certain extent due to the fact that an effective active phase regulation effect is not formed; too high carbon deposition on the catalyst can cause surface coverage of the active phase, resulting in too low activity of the catalyst.

Wherein the yield of naphtha (%) = weight of naphtha/total weight of product 100%

TABLE 2

The embodiment and the comparative example show that the starting method overcomes the problems that in the prior art, on-line sulfuration and passivation steps are needed in the hydrocracking transposed starting process, certain safety and environmental protection risks exist, the selectivity of the catalyst is difficult to improve, and the yield of the target product is low.

The preferred embodiments of the present invention have been described in detail above, but the present invention is not limited thereto. Within the scope of the technical idea of the invention, a number of simple variants of the technical solution of the invention are possible, including combinations of the individual technical features in any other suitable way, which simple variants and combinations should likewise be regarded as being disclosed by the invention, all falling within the scope of protection of the invention.

Claims

1. A method of starting up, the method comprising:

2. The process of claim 1, wherein in step (1), the ammonia-passivated sulfided hydrocracking catalyst is prepared by: the oxidation state hydrocracking catalyst is firstly passivated outside the hydrocracking device and then vulcanized;

And/or, in step (1), the oxidation state hydrocracking catalyst comprises a cracking component, a hydrogenation component and a support;

And/or the cracking component comprises an amorphous acid component and/or a molecular sieve, wherein the amorphous acid component comprises amorphous silica alumina and/or amorphous silica magnesia, and the molecular sieve is at least one selected from a Y-type molecular sieve, a ZSM-5 molecular sieve, a SAPO molecular sieve and an MCM-41 mesoporous molecular sieve;

and/or the hydrogenation component comprises at least one of a group VIII metal and a group VIB metal;

and/or the carrier comprises a refractory porous material selected from at least one of alumina, silica, titania, magnesia, zirconia, and activated carbon;

And/or, in the step (1), the content of the cracking component is 10-60 wt% based on the weight of the oxidized hydrocracking catalyst, the content of the carrier is 30-70 wt%, the content of the group VIII metal calculated as oxide is 1-15 wt%, and the content of the group VIB metal calculated as oxide is 5-30 wt%.

3. The method of claim 2, wherein the passivating process comprises: loading an organic alcohol compound and an organic nitrogen compound on an oxidation state hydrogenation catalyst by adopting an impregnation method in the presence of a solution containing the organic alcohol compound and the nitrogen compound, and then carrying out passivation treatment;

Wherein the organic nitrogen-containing compound and the organic alcohol compound are used in an amount such that the content of nitrogen in the ammonia passivation type oxidation state hydrocracking catalyst is 0.1-8% by weight of the oxidation state hydrocracking catalyst before passivation, and the molar ratio of the organic alcohol compound to the VIII group metal atom in the oxidation state hydrocracking catalyst is 0.1-3.

4. A method according to claim 2 or 3, wherein the vulcanization treatment is wet vulcanization or dry vulcanization;

preferably, the vulcanization treatment includes: under the existence of a vulcanizing agent and hydrogen, vulcanizing the passivated catalyst to obtain an ammonia passivation type vulcanized hydrocracking catalyst;

Preferably, the vulcanizing agent in the wet vulcanization is selected from at least one of carbon disulfide, dimethyl disulfide and polysulfide;

Preferably, the sulfiding agent in the dry sulfiding is selected from hydrogen sulfide and/or hydrogen;

preferably, the conditions of the dry vulcanization treatment include: the heating rate is 5-100 ℃/h, the highest vulcanization temperature is 240-400 ℃, the constant temperature time is 1-15h, the vulcanization pressure is 0.1-10MPa, and the gas agent volume ratio is 50-1000:1, a step of; further preferably, the conditions of the dry vulcanization treatment include: the heating rate is 10-80 ℃/h, the highest vulcanization temperature is 260-380 ℃, the constant temperature time is 2-12h, the vulcanization pressure is 0.3-8MPa, and the gas agent volume ratio is 100-800:1, a step of;

Preferably, the conditions of the wet vulcanization treatment include: the heating rate is 5-100 ℃/h, the highest vulcanization temperature is 240-400 ℃, the constant temperature time is 1-15h, the vulcanization pressure is 0.1-10MPa, the volume airspeed is 0.3-8 hours ^-1, and the hydrogen-oil volume ratio is 100-1000:1, a step of; further preferably, the conditions of the wet vulcanization treatment include: the heating rate is 10-80 ℃/h, the highest vulcanization temperature is 260-380 ℃, the constant temperature time is 2-12h, the vulcanization pressure is 0.3-8MPa, the volume airspeed is 0.5-6 hours ^-1, and the hydrogen-oil volume ratio is 150-800:1.

5. A process according to any one of claims 1 to 3, wherein in step (2) the catalyst bed is heated at a rate of from 5 to 25 ℃/h, preferably from 10 to 25 ℃/h;

Preferably, in the step (2), hydrogen is introduced to raise the temperature of the catalyst bed to 80-120 ℃.

6. A process 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 and hydrocracked diesel, preferably straight run diesel and/or hydrocracked diesel;

And/or the distillation range of the diesel fraction is 210-380 ℃, preferably 210-370 ℃;

and/or the nitrogen content of the diesel fraction is 100-800ppm, preferably 100-500ppm.

7. A 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, preferably at least one of corn oil, soybean oil, peanut oil, rapeseed oil, coconut oil, sunflower oil and cottonseed oil;

and/or the vegetable oil has a distillation range of 230-580 ℃, preferably 230-560 ℃;

and/or the bromine valence of the vegetable oil is 8-20gBr/100mL, preferably 10-15gBr/100mL;

And/or the secondary processing oil is selected from at least one of catalytic gasoline, pyrolysis gasoline, catalytic diesel and coker diesel;

And/or the secondary processing oil has a distillation range of 170-380 ℃, preferably 190-370 ℃;

And/or the secondary processing oil has a density of 0.85-0.98g.cm ^-3, preferably 0.89-0.96g.cm ^-3;

And/or the secondary processing oil has an aromatic hydrocarbon content of 50 to 90 wt.%, preferably 60 to 90 wt.%.

8. A process according to any one of claims 1 to 3, wherein in step (2), the diesel fraction is used in an amount of 200 to 2500 parts by weight, the vegetable oil is used in an amount of 0 to 250 parts by weight, and the secondary processing oil is used in an amount of 0 to 250 parts by weight, relative to 100 parts by weight of the hydrocracking catalyst;

Preferably, the diesel fraction is used in an amount of 300 to 2000 parts by weight, the vegetable oil is used in an amount of 20 to 200 parts by weight, and the secondary processing oil is used in an amount of 10 to 200 parts by weight, compared to 100 parts by weight of the hydrocracking catalyst.

9. A method according to any one of claims 1-3, wherein in step (2), the closed cycle conditions comprise: the pressure is 2-6MPa, and the hydrogen oil volume ratio is 500:1-1300:1, the volume airspeed is 0.3-2.5h ^-1;

Preferably, the pressure is 3-6MPa, and the hydrogen-oil volume ratio is 600:1-1100:1, the volume space velocity is 0.3-1.8h ^-1.

10. A process according to any one of claims 1 to 3 wherein in step (3) the catalyst bed temperature of the previous stage is from 210 to 230 ℃;

And/or, in the step (3), the heat preservation time of the previous stage is 4-10h, and the pressure is 2-6MPa; preferably, the heat preservation time of the previous stage is 4-8 hours, and the pressure is 3-6MPa;

and/or, in the step (3), the temperature rising rate of the previous stage is 5-20 ℃/h, preferably 5-15 ℃/h.

11. A process according to any one of claims 1 to 3, wherein in step (3) the latter stage pressure is 7 to 16MPa, preferably 7 to 14MPa;

And/or, in the step (3), the temperature of the catalyst bed layer in the later stage is 290-320 ℃;

And/or, in the step (3), the temperature rising rate of the catalyst bed layer in the later stage is 5-15 ℃/h, preferably 5-12 ℃/h.

12. A method according to any one of claims 1 to 3, wherein in step (3), the feedstock oil is selected from at least one of straight run wax oil, straight run diesel oil and catalytic diesel oil;

Preferably, the density of the raw oil is 0.85-1g.cm ^-3, the sulfur content is 0.5-1.5wt%, the nitrogen content is 500-5000ppm, and the distillation range is 300-500 ℃.