A method for starting a cracking unit

Abstract

The present invention relates to the technical field of hydrocracking, and discloses a method for starting a cracking unit. The method includes: (1) loading an ammonia-passivated sulfided hydrocracking catalyst into a reaction zone of a hydrocracking unit; (2) in the presence of hydrogen, heating the catalyst bed temperature to 140-180°C, then introducing kerosene fractions and optionally vegetable oil and secondary processed gasoline to wet the catalyst bed, and establishing a closed loop; (3) continuing to heat the catalyst bed temperature described in step (2) to 260-320°C, then replacing the feedstock oil, continuing to heat to the reaction temperature, and switching to production. The method provided by the present invention has a simple start-up process and can effectively improve the yield of BTX in the product.

Description

Technical Field

[0001] The invention relates to the technical field of hydrocracking, and in particular to a method for starting a cracking device. Background Art

[0002] In the face of the market environment of transformation from oil refining to chemical industry and fuel oil to chemical products, it is of great significance to help improve my country's independent supply capacity of aromatics market and reduce its dependence on foreign markets. At the same time, with the continuous development of new energy, more and more electric vehicles and hydrogen fuel vehicles have replaced the original fuel vehicles, resulting in a gradual decline in diesel consumption. my country's crude oil is becoming heavier and heavier. Catalytic cracking units are an important means to achieve light oil products. The aromatic content of the catalytic cracking diesel produced is mostly above 60%. Therefore, through the appropriate processing route, catalytic cracking is a high-quality raw material for the production of BTX. Most of the aromatic components in straight-run diesel and catalytic cracking diesel are mainly dicyclic. If the basic chemical raw material BTX is to be produced, single-ring saturation and ring-opening chain breaking are required. Therefore, diesel hydrocracking is an important route for the production of BTX.

[0003] Highly active hydrocracking catalysts are mainly composed of sulfurized active metals and acidic materials, which play the role of hydrogenation and cracking respectively. In order to give full play to the catalyst or to ensure the safety of the catalyst startup process at the same time, the catalyst must be sulfurized and the acidic material passivated during the startup phase of the hydrocracking unit, in order to improve the hydrogenation activity of the catalyst and temporarily reduce the cracking activity of the catalyst, respectively, to avoid overheating of the bed temperature during the startup process. At present, the ex-situ sulfurization and ex-situ passivation methods are the simplest and most effective catalyst treatment methods. On the one hand, they can effectively reduce the danger of the catalyst during the online sulfurization process, simplify the startup process, and ensure the activity of the catalyst. On the other hand, the ex-situ passivation method ensures the safety of the catalyst startup process. Passivating the acidic material in advance can ensure that the temperature range of the oil inlet during the startup phase of the unit is wider and more flexible.

[0004] Diesel is produced into the basic chemical raw material BTX by hydrocracking, in which the conversion rate or yield is an important indicator. Highly selective catalysts or matching processes can help improve the yield of the product BTX. Catalyst regulation during the start-up process can improve the selectivity of the catalyst to a certain extent, thereby further improving the yield of the target product.

[0005] Patent application CN109777472A discloses a method for hydrogenation start-up. In the catalyst passivation stage, the pre-loaded nitride is used as a passivator to passivate the cracking agent, thereby achieving the purpose of not using a passivator. However, during the start-up process, a sulfiding agent needs to be continuously injected to sulfide the catalyst, and a dedicated sulfiding equipment needs to be installed. In addition, there are problems such as incomplete sulfidation, hydrogen sulfide leakage, long start-up time and high cost during the sulfidation process.

[0006] 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.

[0007] Patent applications CN103059969A and CN103789023A disclose a method for passivating a cracking agent by using ammonia generated by the reaction of high-nitrogen raw materials with hydrogen as a passivating agent during the start-up process, thereby achieving the purpose of reducing the use of anhydrous liquid ammonia. Although it can temporarily inhibit the cracking activity, the operation is relatively complicated and the stability is poor. It is difficult to ensure the passivation effect of the molecular sieve. At the same time, the additional introduction of high-nitrogen raw oil is likely to introduce new impurities into the reaction system, affecting the activity of the catalyst.

[0008] 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.

[0009] Patent application CN103214332A discloses a method for producing light aromatics and high-quality oil products from catalytic cracking diesel. The catalytic cracking diesel is first extracted to obtain an extracted oil rich in polycyclic aromatics and a raffinate oil rich in alkanes, and then the two raw materials are processed separately to obtain the target aromatic product. This method results in that the extraction process can only be carried out alone, and the extracted raw materials cannot be mixed with other raw materials for extraction, which will increase the energy consumption of production.

[0010] Patent application CN104560164A discloses a method for producing a high-octane gasoline component or product BTX fraction by hydro-reforming. Using low-quality diesel fraction as raw material, high-octane gasoline is produced by hydro-reforming. The BTX content in the high-octane gasoline fraction can reach 40% by mass, but it is difficult to directly obtain the raw material for BTX extraction. At the same time, low-quality raw materials require more stringent operating conditions, resulting in a certain degree of loss of aromatics.

[0011] Patent application CN110551525A discloses a method for producing BTX fraction from catalytic cracking diesel. The invention discloses cutting catalytic cracking diesel into light catalytic cracking diesel fraction and heavy catalytic cracking diesel fraction. The light catalytic cracking diesel fraction enters the low-pressure hydrocracking unit, and the heavy catalytic cracking diesel fraction enters the hydrotreating unit before entering the catalytic cracking unit. This method requires fractionation of the raw materials, which will significantly increase the energy consumption of the device. At the same time, the process involves the combination of multiple devices, and the processing process is highly complex. Summary of the invention

[0012] The purpose of the present invention is to overcome the problems of complicated start-up process of hydrocracking unit and low yield of BTX produced by diesel hydrocracking in the prior art, and to provide a start-up method of cracking unit, which has the effect of simple start-up process and improving the yield of BTX in the product to a certain extent.

[0013] In order to achieve the above object, the present invention provides a method for starting up a cracking unit, wherein the method comprises:

[0014] (1) loading an ammonia-passivated sulfided hydrocracking catalyst into a reaction zone of a hydrocracking unit;

[0015] (2) in the presence of hydrogen, raising the temperature of the catalyst bed to 140-180° C., then introducing kerosene fraction and optionally vegetable oil and secondary processed gasoline to wet the catalyst bed, thereby establishing a closed loop;

[0016] (3) The catalyst bed temperature in step (2) is further raised to 260-320° C., and then the crude oil is replaced and the temperature is further raised to the reaction temperature to start production.

[0017] The inventors of the present invention have found in their research that the use of the kerosene fraction of the present invention and optionally vegetable oil and secondary processed gasoline in combination as the start-up oil under the start-up conditions, and the use of the hydrocracking catalyst of the present invention at the same time, can simplify the start-up steps of the hydrocracking unit, and the catalyst can be selectively regulated and controlled during the start-up process by combining the process conditions in the start-up process to improve the selectivity of the catalyst. Using diesel as a raw material can increase the yield of BTX in the product. DETAILED DESCRIPTION

[0018] 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.

[0019] The present invention provides a method for starting up a cracking device, wherein the method comprises:

[0020] (1) loading an ammonia-passivated sulfided hydrocracking catalyst into a reaction zone of a hydrocracking unit;

[0021] (2) in the presence of hydrogen, raising the temperature of the catalyst bed to 140-180° C., then introducing kerosene fraction and optionally vegetable oil and secondary processed gasoline to wet the catalyst bed, thereby establishing a closed loop;

[0022] (3) The catalyst bed temperature in step (2) is further raised to 260-320° C., and then the crude oil is replaced and the temperature is further raised to the hydrocracking reaction temperature to start production.

[0023] The method of the present invention can simplify the start-up steps of the hydrocracking unit. During the start-up process, the catalyst can be selectively regulated by coordinating the process conditions during the start-up process to improve the selectivity of the catalyst. Using diesel as a raw material can increase the yield of BTX in the product.

[0024] In the present invention, there is no particular limitation on the preparation method of the ammonia-passivated sulfurized hydrocracking catalyst in step (1), and the catalyst can be prepared according to the method in patent application CN111729690A. Preferably, in step (1), the ammonia-passivated sulfurized hydrocracking catalyst is prepared by the following method: the hydrocracking catalyst is subjected to a sulfurization treatment and a passivation treatment in sequence outside the hydrocracking unit. In view of the problems of catalyst sulfurization and acidic molecular sieve ammonia passivation during the start-up of the existing hydrocracking catalyst, the present invention provides a catalyst sulfurization process that is completed outside the device, which effectively converts the active metal into a metal sulfide with high catalytic activity, has a good sulfurization effect, avoids the safety and environmental risks caused by the use of a sulfurizing agent at the in-device sulfurization site, and avoids the risk of concentrated decomposition and over-temperature of sulfides that may occur during the in-device activation of the sulfur-loaded catalyst, while reducing the damage of the sulfurizing agent to the human body and the investment in sulfur injection equipment; further, the ammonia passivation process is completed outside the device to effectively avoid the passivation during the start-up process. It can reduce the potential danger of poor chemical reaction or ammonia injection, reduce the use and emission of toxic substances such as ammonia, reduce the waste of passivating agents, save resources, be low-carbon and environmentally friendly, avoid the unstable passivation caused by high-nitrogen oil and the possible introduction of other impurities, and have a high-efficiency and stable passivation effect; further, the addition of low-boiling point organic solvents helps to carry organic nitrogen-containing compounds to better combine with the acidic sites on the molecular sieve, complete the ammonia passivation of the molecular sieve, and at the same time, most of the low-boiling point solvents are removed during the preparation process, avoiding the influence of subsequent solvents on the catalyst activity.

[0025] 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.

[0026] 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.

[0027] 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.

[0028] 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.

[0029] 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.

[0030] In the present invention, there is no particular limitation on the specific means of vulcanization treatment, which may be dry vulcanization or wet vulcanization, and the present invention does not make any particular limitation on this.

[0031] In the present invention, there is no particular limitation on the specific conditions of the sulfurization treatment, which can be prepared according to the method and conditions in patent application CN111729690A, and is not specifically limited in the present invention.

[0032] According to a preferred embodiment of the present invention, in step (1), the passivation treatment comprises: in the presence of an organic solvent, the organic nitrogen-containing compound is loaded onto the sulfided hydrogenation catalyst by an impregnation method, and the solvent is evaporated before passivation treatment is performed to obtain an ammonia-passivated sulfided hydrocracking catalyst.

[0033] According to a preferred embodiment of the present invention, the organic solvent is selected from at least one of hydrocarbon oil, hydrocarbon oil oxygen-containing derivatives and organic carboxylic acid esters.

[0034] According to a preferred embodiment of the present invention, the hydrocarbon oil and hydrocarbon oil oxygen-containing derivatives are each independently selected from at least one of alcohols, ethers and light fraction hydrocarbon oils, and the organic carboxylic acid ester is fatty acid glyceride.

[0035] According to a preferred embodiment of the present invention, the organic solvent is preferably at least one selected from ethanol, propanol, butanediol, ether, cyclohexane, n-heptane, n-decane, methylcyclopentane, naphtha, gasoline, kerosene, diesel, white oil, kerosene and lubricating oil base oil.

[0036] According to a preferred embodiment of the present invention, the organic solvent has 2-35 carbon atoms, preferably 2-15, and more preferably 2-10 carbon atoms.

[0037] 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.

[0038] According to a preferred embodiment of the present invention, the organic nitrogen-containing compound has 1-20 carbon atoms, preferably 2-15 carbon atoms.

[0039] According to a preferred embodiment of the present invention, the organic nitrogen-containing compound and the organic solvent are used in such an amount that the nitrogen content in the ammonia-passivated sulfided hydrocracking catalyst, calculated as an element, accounts for 0.1-3%, preferably 0.5-3%, of the weight of the sulfided hydrocracking catalyst before passivation.

[0040] In the present invention, a conventionally defined impregnation method may be selected, such as saturated impregnation, unsaturated impregnation or supersaturated impregnation. Preferably, the loading may be carried out in at least one of the following ways, such as impregnating the sulfided hydrogenation catalyst into a solution containing an organic nitrogen-containing compound in the presence of an organic solvent; or spraying the solution containing an organic nitrogen-containing compound onto the sulfided hydrogenation catalyst in the presence of an organic solvent.

[0041] In the present invention, there is no particular limitation on the conditions for loading the organic nitrogen-containing compound by the impregnation method. Preferably, the loading temperature is 10-100°C, preferably 20-80°C.

[0042] In the present invention, the evaporation temperature is 10-150° C., the absolute pressure is 0.01-0.5 MPa, the temperature is preferably 20-90° C., the absolute pressure is preferably 0.03-0.3 MPa, and the evaporation is preferably carried out in a rotary evaporator.

[0043] In the present invention, the passivation conditions have a wide selection range. Preferably, the passivation treatment conditions include: temperature of 40-200°C, pressure of 0.01-0.8MPa, and time of 1-8h; further preferably, the passivation treatment conditions include: temperature of 60-180°C, pressure of 0.03-0.6MPa, and time of 2-6h; preferably, the passivation treatment is carried out under a passivation atmosphere; preferably, the passivation atmosphere is selected from at least one of an inert gas, an oxygen atmosphere, and air; preferably, the inert atmosphere is selected from at least one of nitrogen, helium, and argon; preferably, the passivation treatment is carried out in a fixed treatment device; preferably, the passivation treatment is carried out in a non-flowing atmosphere, a naturally flowing atmosphere, or a forced flowing atmosphere.

[0044] In the present invention, hydrogen is introduced into step (2) as a heating medium to increase the temperature of the catalyst bed. In the present invention, there is no particular limitation on the heating rate of the catalyst bed. Preferably, in step (2), the heating rate of the catalyst bed is 10-35°C / h, more preferably 10-30°C / h.

[0045] 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 select it according to actual needs. Preferably, in step (2), the catalyst bed temperature is 150-180° C. The advantage of adopting this preferred embodiment is that a suitable temperature can be used to further remove part of the organic solvent remaining on the catalyst, thereby avoiding the influence on the catalyst activity.

[0046] In the present invention, there is no particular limitation on the type of kerosene fraction in step (2), and those skilled in the art can select it according to actual needs. Preferably, in step (2), the kerosene fraction is selected from at least one of straight-run kerosene, hydrocracked kerosene and Fischer-Tropsch kerosene, preferably straight-run kerosene and / or hydrocracked kerosene.

[0047] According to a preferred embodiment of the present invention, the distillation range of the kerosene fraction is 140-300°C, preferably 160-300°C.

[0048] According to a preferred embodiment of the present invention, the sulfur content in the kerosene fraction is 0.02-0.3wt%, preferably 0.02-0.2wt%.

[0049] By selecting the kerosene fraction according to the above preferred embodiment, the raw material quality for subsequent processing by the device is better, and the dust and impurities carried on the catalyst can be better washed away, and the organic solvent remaining on the catalyst can be dissolved.

[0050] In the present invention, the above-mentioned kerosene fractions and vegetable oils and secondary processed gasoline 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.

[0051] According to a preferred embodiment of the present invention, the distillation range of the vegetable oil is 170-550°C, and the bromine value is 2-15gBr / 100mL; preferably, the distillation range of the vegetable oil is 190-460°C, and the bromine value is 2-10gBr / 100mL.

[0052] Preferably, in step (2), the secondary processed gasoline is selected from at least one of catalytic cracking gasoline, catalytic pyrolysis gasoline, steam pyrolysis gasoline and reforming raffinate.

[0053] According to a preferred embodiment of the present invention, the distillation range of the secondary processed gasoline is 50-205°C and the density is 0.72-0.81 g.cm -3 Preferably, the distillation range of the secondary processed gasoline is 70-190°C and the density is 0.73-0.80g.cm -3 .

[0054] According to a preferred embodiment of the present invention, the olefin content in the secondary processed gasoline is 10-35% by weight, preferably 15-30% by weight.

[0055] According to a preferred embodiment of the present invention, the diene value of the secondary processed gasoline is 3-15 gI / 100g, preferably 5-12 gI / 100g.

[0056] The selection of vegetable oil and secondary processed gasoline according to the above preferred implementation mode has the advantage of regulating the selectivity and activity of the catalyst. The vegetable oil has a long carbon chain and is rich in unsaturated bonds. At a certain temperature, it can form a carbon deposit precursor around the active phase of the catalyst, which has a barrier effect on the growth of the active phase and improves the activity of the catalyst. The secondary processed gasoline contains dienes and olefin substances with high polarity, which can be preferentially adsorbed on the acidic sites and regulate the acidic materials. At the same time, the dienes also have good condensation precursors, which can regulate the morphology and structure of the active phase, and finally work together to improve the activity and selectivity of the catalyst.

[0057] 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 kerosene fraction is 100-1800 parts by weight, and the amount of the vegetable oil is 0-200 parts by weight; the amount of the secondary processed gasoline is 0-200 parts by weight; preferably, compared with 100 parts by weight of the hydrocracking catalyst, the amount of the kerosene fraction is 150-1500 parts by weight, the amount of the vegetable oil is 10-150 parts by weight, and the amount of the secondary processed gasoline is 10-150 parts by weight. The advantage of adopting this preferred embodiment is that through the appropriate ratio, the synergistic effect between the long carbon chain unsaturated molecules in the vegetable oil and the dienes and olefin molecules in the gasoline can be promoted, so that the activity and selectivity of the catalyst can be better adjusted.

[0058] According to a preferred embodiment of the present invention, in step (2), the closed-loop circulation conditions include: a pressure of 6-15 MPa, a hydrogen-to-oil volume ratio of 400:1-1200:1, and a volume space velocity of 0.1-2.5 h -1 Preferably, the pressure is 6-12 MPa, the hydrogen-oil volume ratio is 500:1-1100:1, and the volume space velocity is 0.3-2h -1 The advantage of adopting this preferred embodiment is that under the cooperation of the above-mentioned process conditions, the molecules with certain regulating functions in the start-up oil and the catalyst can be better combined to achieve activity and selectivity regulation during the start-up process. At the same time, the above-mentioned process conditions are relatively close to the subsequent reaction conditions, which is convenient for adjustment to the subsequent normal operating conditions.

[0059] According to a preferred embodiment of the present invention, in step (3), the temperature of the catalyst bed in step (2) is further increased to 260-320° C. The advantages of adopting this preferred embodiment are, on the one hand, preventing the light components in the start-up oil from cracking at too high a temperature and affecting the operation of the subsequent fractionation tower, and on the other hand, better regulating effects on the activity and selectivity of the catalyst in this temperature range.

[0060] According to a preferred embodiment of the present invention, in step (3), the heating rate of the catalyst bed temperature is 5-25°C / h, preferably 10-25°C / h. The advantages of adopting this preferred embodiment are that on the one hand, the heating rate matches the heating furnace load, and on the other hand, at this heating rate, the catalytic strength can be better protected to avoid the influence of too fast heating rate on the strength of the catalyst, and the appropriate heating rate matches the regulating effect of the catalyst activity and selectivity.

[0061] In the present invention, there is no particular limitation on the type of feedstock oil. Preferably, in step (3), the feedstock oil is selected from at least one of straight-run diesel, catalytic diesel, coker diesel and vacuum light fraction.

[0062] According to a preferred embodiment of the present invention, the density of the feedstock oil is 0.79-0.95 g.cm -3 , sulfur content 0.1-1.5wt%, nitrogen content 0.01-0.2wt%, distillation range 200-450℃.

[0063] In the present invention, it is understood that the production conditions, such as pressure, hydrogen-to-oil volume ratio and volume space velocity, are the same or different from the range of conditions for the closed-loop cycle in step (2), and are preferably the same, and the present invention does not impose any specific limitation thereto.

[0064] 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.

[0065] The properties of the start-up oil and feed oil used in the following examples and comparative examples are shown in Table 1.

[0066] Table 1

[0067]

[0068] To illustrate the features of the present invention, the embodiments and comparative examples are all based on the same batch of industrially produced commercial oxidized hydrocracking catalysts with a Ni oxide content of 3.8 wt%, a Mo oxide content of 22.5 wt%, a Y-type molecular sieve content of 50 wt% and the balance being alumina. Dry sulfidation is selected, and a combined gas of hydrogen sulfide and hydrogen is used for sulfidation. The sulfidation treatment is carried out outside the hydrocracking unit. The sulfidation treatment conditions are: H2S volume fraction of 3%, H2 volume fraction of 97%, heating rate of 20°C / h, sulfidation temperature of 340°C, constant temperature time of 8h, sulfidation pressure of 4MPa, and gas-to-agent volume ratio of 400:1, to obtain a sulfided hydrocracking catalyst.

[0069] Example 1

[0070] The hydrocracking unit was filled with hydrocracking catalyst. 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 30°C / h, so that the bed temperature was controlled at 180°C. The start-up kerosene A was introduced into the hydrocracking unit to wet and flush the bed. After the bed was fully wetted and flushed, the start-up kerosene A, vegetable oil C and secondary processed gasoline E were switched to perform closed-loop circulation. The weight ratio of kerosene A to catalyst was 15, the weight ratio of vegetable oil C to catalyst was 1.5, and the weight ratio of secondary processed gasoline E to catalyst was 1.5. The closed-loop circulation was performed with an operating pressure of 8MPa, a hydrogen-to-oil volume ratio of 1000:1, and a volume space velocity of 1.2h. -1 The bed layer was heated at a rate of 25°C / h. When the temperature reached 310°C, the feedstock oil was switched and the reaction temperature was gradually increased through the heating furnace until the product was qualified. At this time, the temperature was 383°C and the reaction began. The operating conditions of the reaction process were a pressure of 8MPa, a hydrogen-to-oil volume ratio of 1000:1, and a volume space velocity of 1.2h -1The product was subjected to simulated distillation. The specific test conditions were as follows: the sample was analyzed by multidimensional gas chromatography SH / T 0558-2016 determination method for 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 by unloading agent. The active phase structure of the catalyst was characterized by HRTEM. The active phase refers to the microstructure with catalytic activity in the catalyst. The specific presentation form is Ni-Mo-S form. The overall layered crystal structure is present. Through high-resolution transmission electron microscopy, it can be observed that the projection of the layered structure is a stripe phase of varying lengths. It is generally believed that the shorter the stripe phase, the higher the catalytic activity. The specific test conditions were as follows: the sample was characterized by Tecnai G2 F20S-TWIN high-resolution transmission electron microscope (HRTEM) produced by FEI Company. The acceleration voltage was 200kV. The size and stacking of MoS2 lamellae in the sample were observed. When preparing the sample, the ground sample was ultrasonically dispersed in a cyclohexane solution, and a small amount of the upper suspension was collected and dropped on a carbon-coated copper mesh.

[0071] The hydrocracking catalyst is an ammonia-passivated sulfided hydrocracking catalyst. After the catalyst is subjected to ex-situ sulfidation, a nitrogen content of 2.35% by weight is loaded. The nitrogen element comes from diethanolamine. The nitrogen-containing compound is loaded onto the sulfided catalyst using ethanol as a solvent. The loading temperature is 60° C., the solvent evaporation temperature is 80° C., the pressure is 0.1 MPa, and the process is completed in a rotary evaporator. The passivation temperature is 150° C., the pressure is 0.1 MPa, the time is 4 hours, and the process is completed in a non-flowing nitrogen atmosphere to prepare an ammonia-passivated sulfided hydrocracking catalyst.

[0072] Example 2

[0073] The hydrocracking catalyst of 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 10°C / h, so that the bed temperature was controlled at 150°C. Start-up kerosene A was introduced into the hydrocracking unit to wet and flush the bed. After the bed was fully wetted and flushed, start-up kerosene A, vegetable oil C and secondary processed gasoline E were switched to perform closed-loop circulation, wherein the weight ratio of kerosene A to the catalyst was 1.5, the weight ratio of vegetable oil C to the catalyst was 0.1, and the weight ratio of secondary processed gasoline E to the catalyst was 0.1. The closed-loop circulation was performed with an operating pressure of 8 MPa, a hydrogen-to-oil volume ratio of 1000:1, and a volume space velocity of 1.2 h -1 The bed layer was heated at a rate of 10°C / h. When the temperature reached 270°C, the feedstock oil was switched and the reaction temperature was gradually increased through the heating furnace until the product was qualified. At this time, the temperature was 382°C and the reaction began. The operating conditions of the reaction process were a pressure of 8MPa, a hydrogen-to-oil volume ratio of 1000:1, and a volume space velocity of 1.2h-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 subjected to catalyst unloading characterization and the active phase structure of the catalyst was subjected to HRTEM characterization analysis.

[0074] Example 3

[0075] The hydrocracking catalyst of 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 20°C / h, so that the bed temperature was controlled at 165°C. The start-up kerosene A was introduced into the hydrocracking unit to wet and flush the bed. After the bed was fully wetted and flushed, the start-up kerosene A, vegetable oil C and secondary processed gasoline E were switched to perform closed-loop circulation, wherein the weight ratio of kerosene A to the catalyst was 10, the weight ratio of vegetable oil C to the catalyst was 0.8, and the weight ratio of secondary processed gasoline E to the catalyst was 0.8. The closed-loop circulation was performed with an operating pressure of 8 MPa, a hydrogen-to-oil volume ratio of 1000:1, and a volume space velocity of 1.2 h -1 The bed layer was heated at a rate of 18°C / h. When the temperature reached 290°C, the feedstock oil was switched and the reaction temperature was gradually increased through the heating furnace until the product was qualified. At this time, the temperature was 380°C and the reaction was officially started. The operating conditions of the reaction process were a pressure of 8MPa, a hydrogen-to-oil volume ratio of 1000:1, and a volume space velocity of 1.2h -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 subjected to catalyst unloading characterization and the active phase structure of the catalyst was subjected to HRTEM characterization analysis.

[0076] Example 4

[0077] The hydrocracking catalyst of 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 20°C / h, so that the bed temperature was controlled at 170°C. The start-up kerosene B was introduced into the hydrocracking unit to wet and flush the bed. After the bed was fully wetted and flushed, the start-up kerosene B, vegetable oil C and secondary processed gasoline E were switched to perform closed-loop circulation, wherein the weight ratio of the kerosene B 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 gasoline E to the catalyst was 1. The closed-loop circulation was performed with an operating pressure of 8 MPa, a hydrogen-to-oil volume ratio of 1000:1, and a volume space velocity of 1.2 h -1 The bed layer was heated at a rate of 20°C / h. When the temperature reached 300°C, the feedstock oil was switched and the reaction temperature was gradually increased through the heating furnace until the product was qualified. At this time, the temperature was 385°C and the reaction began. The operating conditions of the reaction process were a pressure of 8MPa, a hydrogen-to-oil volume ratio of 1000:1, and a volume space velocity of 1.2h-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 subjected to catalyst unloading characterization and the active phase structure of the catalyst was subjected to HRTEM characterization analysis.

[0078] Example 5

[0079] The hydrocracking catalyst of 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 18°C / h, so that the bed temperature was controlled at 160°C. Start-up kerosene A was introduced into the hydrocracking unit to wet and flush the bed. After the bed was fully wetted and flushed, the start-up kerosene A, vegetable oil D and secondary processed gasoline E were switched to perform closed-loop circulation, wherein the weight ratio of kerosene A to the catalyst was 8, the weight ratio of vegetable oil D to the catalyst was 1.2, and the weight ratio of secondary processed gasoline E to the catalyst was 1.2. The closed-loop circulation was performed with an operating pressure of 8 MPa, a hydrogen-to-oil volume ratio of 1000:1, and a volume space velocity of 1.2 h -1 The bed layer was heated at a rate of 15°C / h. When the temperature reached 280°C, the feedstock oil was switched and the reaction temperature was gradually increased through the heating furnace until the product was qualified. At this time, the temperature was 387°C and the reaction began. The operating conditions of the reaction process were a pressure of 8MPa, a hydrogen-to-oil volume ratio of 1000:1, and a volume space velocity of 1.2h -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 subjected to catalyst unloading characterization and the active phase structure of the catalyst was subjected to HRTEM characterization analysis.

[0080] Example 6

[0081] The hydrocracking catalyst of Example 1 was added to 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 15°C / h, so that the bed temperature was controlled at 170°C. Start-up kerosene A was introduced into the hydrocracking unit to wet and flush the bed. After the bed was fully wetted and flushed, start-up kerosene A, vegetable oil C and secondary processed gasoline F were switched to perform closed-loop circulation, wherein the weight ratio of kerosene A to the catalyst was 12, the weight ratio of vegetable oil C to the catalyst was 1, and the weight ratio of secondary processed gasoline F to the catalyst was 1.2. The closed-loop circulation was performed with an operating pressure of 8 MPa, a hydrogen-to-oil volume ratio of 1000:1, and a volume space velocity of 1.2 h -1 The bed layer was heated at a rate of 22°C / h. When the temperature reached 300°C, the feedstock oil was switched and the reaction temperature was gradually increased through the heating furnace until the product was qualified. At this time, the temperature was 388°C and the reaction began. The operating conditions of the reaction process were a pressure of 8MPa, a hydrogen-to-oil volume ratio of 1000:1, and a volume space velocity of 1.2h-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 subjected to catalyst unloading characterization and the active phase structure of the catalyst was subjected to HRTEM characterization analysis.

[0082] Example 7

[0083] The hydrocracking catalyst of 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 33°C / h, so that the bed temperature was controlled at 145°C. Start-up kerosene A was introduced into the hydrocracking unit to wet and flush the bed. After the bed was fully wetted and flushed, start-up kerosene A, vegetable oil C and secondary processed gasoline E were switched to perform closed-loop circulation, wherein the weight ratio of kerosene A to catalyst was 16, the weight ratio of vegetable oil C to catalyst was 1.8, and the weight ratio of secondary processed gasoline E to catalyst was 0.5. The closed-loop circulation was performed with an operating pressure of 8 MPa, a hydrogen-to-oil volume ratio of 1000:1, and a volume space velocity of 1.2 h -1 The bed layer was heated at a rate of 5°C / h. When the temperature reached 260°C, the feedstock oil was switched and the reaction temperature was gradually increased through the heating furnace until the product was qualified. At this time, the temperature was 390°C and the reaction began. The operating conditions of the reaction process were a pressure of 8MPa, a hydrogen-to-oil volume ratio of 1000:1, and a volume space velocity of 1.2h -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 subjected to catalyst unloading characterization and the active phase structure of the catalyst was subjected to HRTEM characterization analysis.

[0084] Example 8

[0085] The hydrocracking catalyst of Example 1 was added to 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 20°C / h, so that the bed temperature was controlled at 170°C. Start-up kerosene A was introduced into the hydrocracking unit to wet and flush the bed. After the bed was fully wetted and flushed, a closed-loop cycle was performed. The operating pressure was 8 MPa, the hydrogen-to-oil volume ratio was 1000:1, and the volume space velocity was 1.2 h -1 The bed layer was heated at a rate of 20°C / h. When the temperature reached 300°C, the feedstock oil was switched and the reaction temperature was gradually increased through the heating furnace until the product was qualified. At this time, the temperature was 390°C and the reaction began. The operating conditions of the reaction process were a pressure of 8MPa, a hydrogen-to-oil volume ratio of 1000:1, and a volume space velocity of 1.2h -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 subjected to catalyst unloading characterization and the active phase structure of the catalyst was subjected to HRTEM characterization analysis.

[0086] Comparative Example 1

[0087] The hydrocracking catalyst of Example 1 was added to 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 15°C / h, so that the bed temperature was controlled at 160°C. Diesel feedstock was introduced into the hydrocracking unit to wet and flush the bed. After the bed was fully wetted and flushed, a closed-loop cycle was performed. The operating pressure was 8 MPa, the hydrogen-to-oil volume ratio was 1000:1, and the volume space velocity was 1.2 h -1 The bed layer was heated at a rate of 15°C / h. When the temperature reached 280°C, it was changed to an open circuit. The reaction temperature was gradually increased by the heating furnace until the product was qualified. At this time, the temperature was 398°C and the reaction began. The operating conditions of the reaction process were a pressure of 8MPa, a hydrogen-to-oil volume ratio of 1000:1, and a volume space velocity of 1.2h -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 subjected to catalyst unloading characterization and the active phase structure of the catalyst was subjected to HRTEM characterization analysis.

[0088] Comparative Example 2

[0089] The hydrocracking catalyst of Example 1 was added to 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 12°C / h, so that the bed temperature was controlled at 160°C. Start-up kerosene A was introduced into the hydrocracking unit to wet and flush the bed. After the bed was fully wetted and flushed, the vegetable oil C was switched to a closed-loop circulation. The operating pressure was 8 MPa, the hydrogen-to-oil volume ratio was 1000:1, and the volume space velocity was 1.2 h -1 The bed layer was heated at a rate of 18°C / h. When the temperature reached 270°C, the feedstock oil was switched and the reaction temperature was gradually increased through the heating furnace until the product was qualified. At this time, the temperature was 396°C and the reaction began. The operating conditions of the reaction process were a pressure of 8MPa, a hydrogen-to-oil volume ratio of 1000:1, and a volume space velocity of 1.2h -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 subjected to catalyst unloading characterization and the active phase structure of the catalyst was subjected to HRTEM characterization analysis.

[0090] Comparative Example 3

[0091] The hydrocracking catalyst of Example 1 was added to 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 20°C / h, so that the bed temperature was controlled at 160°C. Start-up kerosene A was introduced into the hydrocracking unit to wet and flush the bed. After the bed was fully wetted and flushed, secondary processed gasoline E was introduced for closed-loop circulation. The operating pressure was 8 MPa, the hydrogen-to-oil volume ratio was 1000:1, and the volume space velocity was 1.2 h -1 The bed layer was heated at a rate of 25°C / h. When the temperature reached 275°C, the feedstock oil was switched and the reaction temperature was gradually increased through the heating furnace until the product was qualified. At this time, the temperature was 396°C and the reaction began. The operating conditions of the reaction process were a pressure of 8MPa, a hydrogen-to-oil volume ratio of 1000:1, and a volume space velocity of 1.2h -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 subjected to catalyst unloading characterization and the active phase structure of the catalyst was subjected to HRTEM characterization analysis.

[0092] Comparative Example 4

[0093] The hydrocracking catalyst of 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 180°C. Start-up kerosene A was introduced into the hydrocracking unit to wet and flush the bed. After the bed was fully wetted and flushed, the start-up kerosene A and the secondary processed gasoline E were switched to perform a closed-loop circulation, wherein the weight ratio of the kerosene A to the catalyst was 10, and the weight ratio of the secondary processed gasoline E to the catalyst was 1. The closed-loop circulation was performed with an operating pressure of 8 MPa, a hydrogen-to-oil volume ratio of 1000:1, and a volume space velocity of 1.2 h -1 The bed layer was heated at a rate of 15°C / h. When the temperature reached 300°C, the feedstock oil was switched and the reaction temperature was gradually increased through the heating furnace until the product was qualified. At this time, the temperature was 394°C and the reaction began. The operating conditions of the reaction process were a pressure of 8MPa, a hydrogen-to-oil volume ratio of 1000:1, and a volume space velocity of 1.2h -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 subjected to catalyst unloading characterization and the active phase structure of the catalyst was subjected to HRTEM characterization analysis.

[0094] Comparative Example 5

[0095] The hydrocracking catalyst of 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 15°C / h, so that the bed temperature was controlled at 170°C. Start-up kerosene A was introduced into the hydrocracking unit to wet and flush the bed. After the bed was fully wetted and flushed, the vegetable oil C and the secondary processed gasoline E were switched to perform a closed-loop circulation, wherein the weight ratio of the vegetable oil C to the catalyst was 1.5, and the weight ratio of the secondary processed gasoline E to the catalyst was 1.5. The closed-loop circulation was performed with an operating pressure of 8 MPa, a hydrogen-to-oil volume ratio of 1000:1, and a volume space velocity of 1.2 h -1 The bed layer was heated at a rate of 10°C / h. When the temperature reached 310°C, the feedstock oil was switched and the reaction temperature was gradually increased through the heating furnace until the product was qualified. At this time, the temperature was 395°C and the reaction began. The operating conditions of the reaction process were a pressure of 8MPa, a hydrogen-to-oil volume ratio of 1000:1, and a volume space velocity of 1.2h -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 subjected to catalyst unloading characterization and the active phase structure of the catalyst was subjected to HRTEM characterization analysis.

[0096] Comparative Example 6

[0097] The hydrocracking catalyst of 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 5°C / h, so that the bed temperature was controlled at 120°C. Start-up kerosene A was introduced into the hydrocracking unit to wet and flush the bed. After the bed was fully wetted and flushed, the start-up kerosene A, vegetable oil C and secondary processed gasoline E were switched to perform closed-loop circulation, wherein the weight ratio of kerosene A to the catalyst was 20, the weight ratio of vegetable oil C to the catalyst was 0.05, and the weight ratio of secondary processed gasoline E to the catalyst was 5. The closed-loop circulation was performed with an operating pressure of 8 MPa, a hydrogen-to-oil volume ratio of 1000:1, and a volume space velocity of 1.2 h -1 The bed layer was heated at a rate of 30°C / h. When the temperature reached 330°C, the feedstock oil was switched and the reaction temperature was gradually increased through the heating furnace until the product was qualified. At this time, the temperature was 398°C and the reaction began. The operating conditions of the reaction process were a pressure of 8MPa, a hydrogen-to-oil volume ratio of 1000:1, and a volume space velocity of 1.2h -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 subjected to catalyst unloading characterization and the active phase structure of the catalyst was subjected to HRTEM characterization analysis.

[0098] The effects of the above embodiments and comparative examples are listed in Table 2 for comparison, wherein, by analyzing the product, each fraction segment reaches the corresponding product index (the initial distillation point -150°C fraction segment is the BTX extraction raw material, and the sulfur content is less than 1ppm and the nitrogen content is less than 1ppm). Under the condition that the product index is qualified, the lower the reaction temperature, the higher the activity of the catalyst; the higher the yield of BTX, the better the selectivity of the catalyst; the shorter the length of the catalyst active phase, the higher the activity potential of the catalyst.

[0099] Wherein, the yield of BTX (%) = weight of BTX / total weight of product*100%.

[0100] The length of the stripe of the active phase of the catalyst is measured by high-resolution transmission electron microscopy. It is necessary to count the length of the stripe sample, take 20 photos, and count the length of the stripe phase on the photos. The average length of the MoS2 flakes Calculated according to the formula.

[0101] Average length of platelets (Unit: nm):

[0102]

[0103] Where: l i is the length of the i-th MoS2 platelet stripe phase; n is the total number of platelets in the statistical area.

[0104] Table 2

[0105]

[0106] It can be seen from the above embodiments and comparative examples that the start-up method can directly switch the start-up oil during the start-up process to start the operation. At the same time, the selectivity of the catalyst is regulated by means of the component characteristics of the vegetable oil and the secondary processed gasoline in the start-up oil. Using diesel as a raw material can effectively increase the yield of BTX in the product. At the same time, the reaction temperature is reduced to a certain extent, which reserves more room for subsequent device temperature increase and provides a guarantee for the long-term operation of the device.

[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 cracking unit, characterized in that: The method includes: (1) Loading an ammonia-passivated sulfided hydrocracking catalyst into the reaction zone of a hydrocracking unit; (2) In the presence of hydrogen, the catalyst bed temperature is raised to 140-180°C, and then kerosene fractions, vegetable oil and secondary processed gasoline are introduced to wet the catalyst bed to establish a closed loop; (3) The catalyst bed temperature in step (2) is further raised to 260-320°C, and then the crude oil is replaced and the temperature is further raised to the reaction temperature to start production; The olefin content in the secondary processed gasoline is 10-35% by weight; The diene value in the secondary processed gasoline is 3-15gI / 100g; Compared to 100 parts by weight of the hydrocracking catalyst, the amount of the kerosene fraction is 150-1500 parts by weight, the amount of the vegetable oil is 10-150 parts by weight, and the amount of the secondary processed gasoline is 10-150 parts by weight; The feedstock oil is selected from at least one of straight-run diesel, catalytic diesel, coker diesel and vacuum light distillate; The distillation range of the vegetable oil is 170-500° C., and the bromine value is 2-15 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: subjecting the hydrocracking catalyst to a sulfiding treatment and a passivation treatment in sequence outside a hydrocracking unit.

3. The method according to claim 2, wherein: In step (1), the passivation treatment comprises: in the presence of an organic solvent, the organic nitrogen-containing compound is loaded onto the sulfided hydrogenation catalyst by an impregnation method, and the passivation treatment is performed to obtain an ammonia-passivated sulfided hydrocracking catalyst; The organic solvent is hydrocarbon oil and / or hydrocarbon oil oxygen-containing derivative; The organic solvent has a carbon number of 2-35; The organic nitrogen-containing compound is selected from at least one of alkylamine compounds, arylamine compounds, amide compounds, alcoholamine compounds and polyamine compounds; The organic nitrogen-containing compound has 1 to 20 carbon atoms; The organic nitrogen-containing compound and the organic solvent are used in such an amount that the nitrogen content in the ammonia-passivated sulfided hydrocracking catalyst, calculated as an element, accounts for 0.1-3% of the weight of the sulfided hydrocracking catalyst before passivation.

4. The method according to claim 3, wherein: The hydrocarbon oil is at least one selected from light fraction hydrocarbon oils.

5. The method according to claim 3, wherein: The organic solvent is an organic carboxylate.

6. The method according to claim 5, wherein: The organic carboxylic acid ester is fatty acid glyceride.

7. The method according to claim 3, wherein: The organic solvent has 2-15 carbon atoms.

8. The method according to claim 7, wherein: The organic solvent has 2-10 carbon atoms.

9. The method according to claim 3, wherein: The organic nitrogen-containing compound is an aniline compound and / or a methylaniline compound.

10. The method according to claim 3, wherein: The organic nitrogen-containing compound is an alkylamine compound and / or an alcoholamine compound.

11. The method according to claim 3, wherein: The organic nitrogen-containing compound has 2-15 carbon atoms.

12. The method according to claim 3, wherein: The organic nitrogen-containing compound and the organic solvent are used in such an amount that the nitrogen content in the ammonia-passivated sulfided hydrocracking catalyst, calculated as an element, accounts for 0.5-3% of the weight of the sulfided hydrocracking catalyst before passivation.

13. The method according to any one of claims 1 to 3, wherein: In step (2), the heating rate of the catalyst bed is 10-35°C / h.

14. The method according to claim 13, wherein: In step (2), the heating rate of the catalyst bed is 10-30°C / h.

15. The method according to any one of claims 1 to 3, wherein: In step (2), the catalyst bed temperature is 150-180°C.

16. The method according to any one of claims 1 to 3, wherein: In step (2), the kerosene fraction is selected from at least one of straight-run kerosene, hydrocracked kerosene and Fischer-Tropsch kerosene; and / or, the distillation range of the kerosene fraction is 140-300° C.; And / or, the sulfur content in the kerosene fraction is 0.02-0.3wt%.

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

18. The method according to claim 16, wherein: The distillation range of the kerosene fraction is 160-300°C.

19. The method according to claim 16, wherein: The sulfur content in the kerosene fraction is 0.02-0.2wt%.

20. 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 secondary processed gasoline is selected from at least one of catalytic cracking gasoline, catalytic pyrolysis gasoline, steam pyrolysis gasoline and reforming raffinate; And / or, the density of the secondary processed gasoline is 0.72-0.81 g.cm -3 , distillation range is 50-205℃; and / or, the olefin content in the secondary processed gasoline is 15-30% by weight; And / or, the diene value of the secondary processed gasoline is 5-12 gI / 100g.

21. The method according to claim 20, 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.

22. The method according to any one of claims 1 to 3, wherein: The distillation range of the vegetable oil is 190-460° C., and the bromine value is 2-10 gBr / 100 mL.

23. The method according to claim 20, wherein: The density of the secondary processed gasoline is 0.73-0.8 g.cm -3 , the distillation range is 70-190℃.

24. The method according to any one of claims 1 to 3, wherein: In step (2), the closed-loop circulation conditions include: pressure of 6-15 MPa, hydrogen-to-oil volume ratio of 400:1-1200:1, and volume space velocity of 0.1-2.5 h -1 .

25. The method according to claim 24, wherein: In step (2), the closed-loop circulation conditions include: pressure of 6-12 MPa, hydrogen-oil volume ratio of 500:1-1100:1, volume space velocity of 0.3-2 h -1 .

26. The method according to any one of claims 1 to 3, wherein: In step (3), the temperature of the catalyst bed in step (2) is further increased to 270-310°C.

27. The method according to claim 26, wherein: In step (3), the heating rate of the catalyst bed temperature is 5-25°C / h.

28. The method according to claim 27, wherein: In step (3), the heating rate of the catalyst bed temperature is 10-25°C / h.

29. The method according to any one of claims 1 to 3, wherein: The density of the raw oil is 0.79-0.95g.cm -3 , sulfur content 0.1-1.5wt%, nitrogen content 0.01-0.2wt%, distillation range 200-450℃.

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

31. The method according to claim 30, 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.

32. The method of claim 30, wherein: The hydrogenation component includes at least one of a Group VIII metal and a Group VIB metal.

33. The method of claim 30, 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.

34. The method of claim 32, wherein: 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.

35. The method of claim 3, wherein: The oxygen-containing derivative of hydrocarbon oil is selected from at least one of alcohols and ethers.

Images