A hydrocracking method for coal-to-liquid

Abstract

The present invention discloses a coal-to-liquid hydrocracking method. In the method, a hydrorefining reaction zone is loaded with a graphite-containing hydrorefining catalyst, and a hydrocracking reaction zone is loaded with a graphite-containing hydrocracking catalyst. The startup process includes: after the catalyst loading and the establishment of an inert gas circulation system are completed, the temperature of each catalyst bed is raised to 440-500°C under an inert gas environment. When the total concentration of CO and CO₂ in the circulating gas reaches 0.6% to 2.0% by volume, the inlet temperature of each reactor is reduced at a rate of ≯20°C / h to 150-190°C. The device is then subjected to hydrogen replacement and sulfurization. After sulfurization is complete, the feedstock is switched to introduce crude oil and resume normal production. The startup method of the present invention can maintain good activity and stability of each catalyst, extending the operating cycle of the device.

Description

Technical Field

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

[0002] Directly using coal as a fuel for heating or power generation is inefficient and economically unprofitable. However, if coal resources can be liquefied directly or indirectly into oil products, and then finely processed and separated through a hydrogenation process, high-value-added products can be obtained.

[0003] Whether produced from direct or indirect coal liquefaction, the oil products often suffer from low sulfur and nitrogen content and poor catalyst stability. At the end of their operation, all coal-to-liquid hydrocracking units often experience issues such as high olefin content in the product, poor stability, and catalyst deactivation due to a reduced temperature rise. Therefore, the development of highly stable hydrorefining and cracking catalysts is crucial to addressing the operational lifecycle issues of coal-to-liquid hydrocracking units.

[0004] CN1854265A discloses a method for producing more middle distillates from Fischer-Tropsch oil. This method involves contacting the full fraction of the Fischer-Tropsch oil with hydrogen and a hydrotreating catalyst. The hydrotreating reaction stream is separated to produce middle distillates, naphtha, and heavy oil. Subsequently, the heavy oil is contacted with hydrogen and an isomerized hydrocracking catalyst, and the hydrocracking reaction effluent is separated to produce a middle distillate product. This method achieves a diesel product yield of up to 80% by weight and a cetane number of approximately 80, making it a high-quality diesel blending component. However, this method does not address catalyst sulfur loss, resulting in a decrease in diesel quality, particularly stability, at the end of operation.

[0005] CN1854264A discloses an integrated Fischer-Tropsch oil hydro-upgrading method. This method aims to maximize the production of high-quality middle distillate oil products. First, the full fraction of the Fischer-Tropsch oil and hydrogen are contacted with a hydrotreating catalyst. The hydrotreating reaction effluent is mixed with the hydrocracking reaction stream and separated to produce middle distillate oil products, naphtha, and tail oil. The tail oil is then mixed with hydrogen and circulated to an isomerization cracking reactor for contact with a hydroisomerization cracking catalyst. This method can produce approximately 80% middle distillate oil, with high quality. However, the catalyst also suffers from poor late-stage stability, making it difficult to maintain the operating cycle of the device. Summary of the Invention

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

[0007] The present invention provides a coal-to-liquid hydrocracking method, wherein the method adopts a hydrorefining reaction zone and a hydrocracking reaction zone, wherein the hydrorefining reaction zone is filled with a hydrorefining catalyst containing graphite, and the content of graphite is 2% to 10%, preferably 2% to 6%, based on the mass of the hydrorefining catalyst containing graphite. The hydrocracking reaction zone is filled with a hydrocracking catalyst containing graphite, and the content of graphite is 2% to 10%, preferably 2% to 6%, based on the mass of the hydrocracking catalyst containing graphite. The method includes: after completing the catalyst loading and the establishment of the inert gas circulation system in the coal-to-liquid hydrocracking unit in sequence, raising the temperature of each catalyst bed to 440-500°C under an inert gas environment; when it is detected that the total concentration of CO and CO2 in the circulating gas reaches 0.6v%-2.0v%, reducing the inlet temperature of each reactor to 150-190°C at a rate of ≯20°C / h (preferably 5-20°C / h); then hydrogen replacement and sulfurization are carried out on the device; after the sulfurization is completed, the feed is switched to introduce raw oil and normal production is resumed.

[0008] Furthermore, the graphite-containing hydrorefining catalyst includes a carrier and an active metal. The carrier of the hydrorefining catalyst includes alumina and graphite, and may also contain a binder. The active metals are Group VIII and Group VIB metals, wherein Group VIII metals are preferably Co and / or Ni, and Group VIB metals are preferably W and / or Mo. The catalyst is preferably prepared as follows: first, the carrier component is uniformly mixed with water and graphite, a binder is added, and the mixture is thoroughly rolled and formed. The catalyst carrier is then dried to obtain a catalyst carrier, which is then impregnated with a solution containing the active metal component. The catalyst carrier is then dried and calcined to obtain a hydrorefining catalyst. The impregnation method may be an equal volume impregnation method, an excess volume impregnation method, or a steam impregnation method. The binder that may be added during the forming process may be a conventional binder, such as at least one of citric acid, butanediol, and glycerol. Conventional forming aids, such as peptizing acid, may also be added during the forming process. The graphite is one of natural flake graphite, earthy graphite and artificial graphite, with a particle size of 10 to 50 μm. The drying conditions for preparing the carrier are as follows: the drying temperature is 40 to 100° C., preferably 50 to 80° C. The drying time is generally 2 to 24 hours, preferably 4 to 12 hours. The drying conditions after impregnation are as follows: the drying temperature is 100 to 200° C., preferably 120 to 150° C., and the drying time is 2 to 24 hours, preferably 4 to 12 hours. The calcination conditions are as follows: under an inert gas environment, the calcination temperature is 250 to 350° C., and the calcination time is 2 to 6 hours. Based on the weight of the catalyst, the content of alumina is 15% to 65%, the content of graphite is 2% to 10%, the content of binder is less than 8%, the content of active metal calculated as oxide is 12% to 50%, of which the content of Group VIII metal calculated as oxide is 1.5% to 10%, and the content of Group VIB metal calculated as oxide is 10% to 40%.

[0009] Furthermore, the graphite-containing hydrocracking catalyst comprises an active metal component and a support, wherein the active metal component is a Group VIII or Group VIB metal, and the support comprises Beta molecular sieve, alumina, and graphite. A binder component may also be included. The Group VIII metal is preferably Co and / or Ni, and the Group VIB metal is preferably W and / or Mo. The graphite-containing hydrocracking catalyst is preferably prepared by the following method: Beta molecular sieve, alumina, water, and graphite are uniformly mixed, a binder is added, and the mixture is thoroughly rolled and formed, followed by drying to obtain a catalyst support; the catalyst support is impregnated with a solution containing the active metal component, and after impregnation, the catalyst support is dried and calcined under an inert gas atmosphere to obtain the hydrocracking catalyst. The impregnation method may include equal volume impregnation, excess volume impregnation, or stepwise impregnation. The binder may be at least one of citric acid, butanediol, and glycerol. Conventional molding aids, such as peptizing acid, may be added during the molding process. The graphite is selected from the group consisting of natural flake graphite, earthy graphite, and artificial graphite, with a particle size of 10 to 50 μm. The drying conditions for preparing the carrier are as follows: a drying temperature of 40 to 100° C., preferably 50 to 80° C. The drying time is generally 2 to 24 hours, preferably 4 to 12 hours. The drying conditions after impregnation are as follows: a drying temperature of 100 to 200° C., preferably 120 to 150° C., and a drying time of 2 to 24 hours, preferably 4 to 12 hours. The calcination conditions are as follows: a calcination temperature of 250 to 350° C., and a calcination time of 2 to 6 hours under an inert gas atmosphere. Based on the mass of the hydrocracking catalyst containing the graphite, the carrier content is 65% to 90%, and the content of the active metal in terms of oxide is 10% to 35%. Based on the mass of the hydrocracking catalyst containing graphite, the molecular sieve content is 5% to 10%, the alumina content is 20% to 60%, the graphite content is 2% to 10%, and the binder content is less than 10%. Based on the mass of the hydrocracking catalyst containing graphite, the content of the Group VIII metal as oxide is 1% to 8%, and the content of the Group VIB metal as oxide is 14% to 32%.

[0010] Furthermore, the process of establishing the inert gas circulation system is carried out using conventional methods in the art and is not particularly limited in the present invention. After the inert gas circulation system is established, the high-pressure portion is controlled to the design pressure, generally 4.0 to 18.0 MPa.

[0011] Furthermore, the hydrogen replacement is completed when the hydrogen volume fraction in the circulating hydrogen reaches 92% to 99%. The hydrogen replacement process can introduce new hydrogen into the circulating gas, and the introduction device at the inlet of the hydrorefining reaction zone replaces the gas in the system until the hydrogen volume fraction in the circulating hydrogen reaches 92% to 99%.

[0012] Furthermore, during the hydrogen replacement, the inlet temperatures of the hydrofining reaction zone and the hydrocracking reaction zone may be lowered to 140-190°C.

[0013] Furthermore, the vulcanization is preferably performed by wet vulcanization.

[0014] Furthermore, the wet sulfurization includes a first constant temperature sulfurization and a second constant temperature sulfurization. The first constant temperature sulfurization is to increase the inlet temperature of the hydrofining reaction zone and the hydrocracking reaction zone to a first constant temperature of 240-260°C at a rate of ≯10°C / h, preferably 2-10°C / h, and then to perform constant temperature sulfurization for 6-10 hours at the first constant temperature. During the sulfurization stage, the mass concentration of hydrogen sulfide in the circulating hydrogen is controlled to be maintained at 1000-6000 ppm. The second constant temperature sulfurization process is as follows: the inlet temperature of the hydrofining reaction zone and the hydrocracking reaction zone is increased to a second constant temperature of 340-380°C, preferably 350-380°C, at a rate of ≯6°C / h (preferably 2-6°C / h), and to perform sulfurization for 2-4 hours at the second constant temperature. During the sulfurization stage, the mass concentration of hydrogen sulfide in the circulating hydrogen is controlled to be maintained at 5000-15000 ppm.

[0015] Furthermore, the wet sulfiding process allows hydrogen sulfide to penetrate the catalyst bed before entering the sulfiding stage. The process is as follows: low-nitrogen oil is introduced from the inlet of the hydrorefining reaction zone to wet the catalyst, a closed-loop circulation of the low-nitrogen oil is established, and then a sulfiding agent is gradually injected into the low-nitrogen oil. The temperature is gradually increased and the catalyst bed temperature in the hydrorefining reaction zone and the hydrocracking reaction zone is controlled not to exceed 230°C until hydrogen sulfide penetrates the entire catalyst bed. Preferably, the conditions for hydrogen sulfide to penetrate the entire catalyst bed are as follows: the mass concentration of hydrogen sulfide in the circulating hydrogen is above 1000 ppm, and more preferably 1000 to 10000 ppm.

[0016] Furthermore, after the first constant temperature sulfidation is completed, a passivating agent is injected at the inlet of the hydrocracking reaction zone, and the injection rate of the sulfiding agent is adjusted so that the hydrogen sulfide concentration in the circulating hydrogen reaches 5000-15000 ppm. The ammonia content in the acidic water discharged from the bottom of the high-pressure separator is monitored. Before ammonia breakthrough, the maximum temperature of the catalyst bed in the hydrocracking reaction zone is controlled to be ≤270°C. After ammonia breakthrough, the passivating agent injection rate is adjusted to maintain the ammonium ion mass concentration in the acidic water discharged from the bottom of the high-pressure separator at 1000-8000 ppm. The addition of the passivating agent is stopped until the sulfidation process is completed. Furthermore, the condition for ammonia breakthrough is: the ammonium ion mass concentration in the acidic water discharged from the bottom of the high-pressure separator reaches above 6000 ppm, preferably 8000-15000 ppm.

[0017] Furthermore, after the second constant temperature vulcanization is completed, the vulcanization process is completed, and the raw oil is switched to normal production.

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

[0019] Furthermore, the passivating agent is selected from anhydrous liquid ammonia and / or an organic passivating agent. The organic passivating agent is one or more of cyclohexylamine, aniline, and n-butylamine. Liquid ammonia can be injected at the inlet of the cracking reactor using a high-pressure ammonia injection pump, or the organic passivating agent can be injected into the raw oil tank or feed pipeline.

[0020] Furthermore, the low nitrogen oil is generally atmospheric distillate oil, and the density is generally 0.76 to 0.86 g / cm 3 The initial distillation point is generally 100-250°C, the final distillation point is generally 280-408°C, and the nitrogen mass content is 100-800ppm.

[0021] Furthermore, after the sulfurization process is completed, the process of switching to normal production of the raw oil is as follows: after the sulfurization is completed, the inlet temperature of the hydrorefining reaction zone is reduced to 260-300°C, and then the inlet temperature of the hydrorefining reaction zone and the cracking reaction zone is increased at a rate of 2-6°C / h until the product quality is qualified. Furthermore, the coal-to-liquid hydrocracking method includes: the coal-to-liquid feedstock is mixed with hydrogen and then enters the hydrorefining reaction zone, where deolefination and deoxygenation reactions occur, the hydrorefining reaction effluent enters the first-stage separation system to obtain light components and heavy components, the heavy components enter the hydrocracking reaction zone to undergo isomerization chain scission reaction, and the hydrocracking reaction effluent enters the second-stage separation system to obtain light naphtha, heavy naphtha, diesel and wax oil products.

[0022] Furthermore, the cut point between the light components and the heavy components separated by the first stage separation system is 165-210°C.

[0023] Furthermore, the feedstock oil is selected from coal-to-liquids, and the initial boiling point of the feedstock oil is 20-100°C, the final boiling point is 300-400°C, the sulfur content is ≤50ppm, and the nitrogen content is ≤20ppm. When the iron content in the coal-to-liquid feedstock exceeds the standard, a protective agent should be added to the top of the upper bed of the hydrorefining catalyst and the hydrocracking catalyst to reduce the pressure drop of the device and extend the operating cycle. The protective agent can be one or more of FZC-105, FZC-106, FBN-03B04, and FBN-03B05 developed by Sinopec Fushun Petrochemical Research Institute.

[0024] Furthermore, during normal production, the reaction conditions in the hydrofining reaction zone are as follows: reaction pressure 6.0-15.0 MPa, preferably 7.0-12.0 MPa; average reaction temperature 270-340°C, preferably 280-320°C; hydrogen-to-oil volume ratio 400:1-2200:1, preferably 600:1-1400:1; liquid hourly volume space velocity 0.5-4.0 h -1 .

[0025] Furthermore, during normal production, the reaction conditions of the hydrocracking reaction zone are generally as follows: reaction pressure of 6.0 to 15.0 MPa, preferably 7.0 to 12.0 MPa; average reaction temperature of 280 to 400°C, preferably 300 to 350°C; hydrogen to oil volume ratio of 400:1 to 2200:1, preferably 800:1 to 1600:1; liquid hourly space velocity of 0.5 to 5.0 h -1 , preferably 1.0~4.0h -1 .

[0026] Compared with the prior art, the method of the present invention has the following beneficial effects:

[0027] 1. The active metals of hydrogenation catalysts are divided into four-coordinate and six-coordinate. The former has a strong interaction with the carrier and good stability, but weak hydrogenation performance; the latter has a weak interaction with the carrier, strong hydrogenation performance, but poor stability. Coal-to-oil raw materials have low sulfur and nitrogen content. The catalyst reacts at high temperature for a long time, which is prone to sulfur loss problems and the device operation time is short. Therefore, for coal-to-oil hydrogenation units, the catalyst used should have good high-temperature stability and anti-sulfurization ability. During the start-up process of the coal-to-oil hydrogenation unit, the present invention dopes part of the graphite into the hydrorefining catalyst and the hydrocracking catalyst. During the start-up process, the six-coordinate active metals with high hydrogenation activity but poor stability in the catalyst are moderately reduced to four-coordinate active metals with good stability, and the sulfurization conditions are adjusted, especially the secondary constant temperature sulfurization temperature is increased, so as to increase the sulfurization degree of the four-coordinate active phase, thereby improving the stability of the catalyst while ensuring the activity of the catalyst.

[0028] 2. When the method of the present invention is used in the coal-to-liquid hydrocracking process, the diesel product obtained is of good quality and has low hydrogen consumption. DETAILED DESCRIPTION

[0029] The coal-to-liquid hydrocracking method provided by the present invention will be further described below with reference to the following examples, but the present invention is not limited thereto.

[0030] Table 1 Properties of low nitrogen oil and feedstock oil

[0031]

[0032]

[0033] Table 2 Industrial catalysts

[0034] Industrial agents FF-66 FC-14 FC-76 Physical and chemical properties Average pore size / nm 5.0nm 4.8nm 5.6nm <![CDATA[Pore volume / mL·g -1 > 0.27 0.33 0.35 <![CDATA[Specific surface area / m 2 ·g -1 > 185 310 192 shape Shamrock Shamrock cylindrical bars <![CDATA[Loading heap ratio, g / cm 3 > 0.72 0.84 0.72

[0035] Table 3 Evaluation conditions

[0036] System pressure, MPa 8.0 <![CDATA[Volumetric hourly space velocity during refining of liquid, h -1 > 2.0 <![CDATA[The hourly space velocity of the cracking agent liquid, h -1 > 2.0 Temperature rise of hydrotreating reactor / hydrocracking reactor, ℃ 55 / 40 Hydrogen-to-oil ratio at the inlet of hydrotreating reactor / hydrocracking reactor 800:1 / 1200:1 Running time, h 2500

[0037] The coal-to-liquid hydrocracking process used in the following examples and comparative examples employs two reactors, wherein the first reactor is a hydrorefining reactor loaded with a hydrorefining catalyst, and the second reactor is a hydrocracking reactor loaded with a hydrocracking catalyst. The process comprises: a coal-to-liquid feedstock is mixed with hydrogen and then enters the hydrorefining reactor, where deolefination and deoxygenation reactions occur. The effluent from the hydrorefining reaction enters a first separation system to obtain light components and heavy components, wherein the cut point between the light components and the heavy components is 165°C. The heavy components enter the hydrocracking reactor for isomerization chain scission reaction. The effluent from the hydrocracking reaction enters a second separation system to obtain light naphtha, heavy naphtha, diesel, and wax oil products.

[0038] The properties of the feedstock oils used in the following Examples and Comparative Examples are shown in Table 1. The Examples were loaded with the catalyst prepared by the method of the present invention, while the Comparative Examples used an FF-66 hydrorefining catalyst in the refining reactor and an FC-14 or FC-76 hydrocracking catalyst in the cracking reactor. The distillation range of the light naphtha in both the Examples and Comparative Examples was <65°C, the distillation range of the heavy naphtha was 65-165°C, the distillation range of the diesel was 165-360°C, and the distillation range of the qualified wax was >360°C.

[0039] Example 1

[0040] The preparation methods of the hydrorefining catalyst and the hydrocracking catalyst involved in this embodiment are as follows:

[0041] Hydrorefining Catalyst: Alumina was mixed with water, nitric acid, and graphite (average particle size 15 μm), and citric acid was added as a binder. The mixture was extruded and dried at 50°C for 10 hours to obtain a shaped carrier. The cobalt and molybdenum active components were impregnated using ammonium molybdate hexahydrate and cobalt nitrate hexahydrate as the molybdenum source via excess volume co-impregnation. The mixture was dried at 120°C for 6 hours and calcined at 300°C under a nitrogen atmosphere for 4 hours to obtain the finished hydrorefining catalyst F-1. The prepared catalyst had an alumina content of 59%, a graphite content of 5%, a cobalt content calculated as CoO of 5.0% by weight, and a molybdenum content calculated as MoO3 of 31.0% by weight, based on the mass of the catalyst.

[0042] Hydrocracking catalyst: Beta molecular sieve, alumina, water, nitric acid, and graphite (average particle size 15 μm) were mixed, and citric acid was added as a binder. The mixture was extruded and dried at 55°C for 12 hours to obtain a shaped support. Molybdenum and nickel active metals were impregnated separately using nickel nitrate hexahydrate as the nickel source and ammonium molybdate hexahydrate as the molybdenum source using an equal volume impregnation method. The mixture was dried at 120°C for 10 hours and calcined at 280°C under a nitrogen atmosphere for 5 hours to obtain the finished hydrocracking catalyst C-1. The catalyst contained 6% beta molecular sieve, 65% alumina, 4% graphite, 5.0% nickel as NiO, and 20.0% molybdenum as MoO3, based on the catalyst mass.

[0043] The hydrotreating catalyst and hydrocracking catalyst prepared above were respectively loaded into the hydrotreating reactor and the hydrocracking reactor, and sulfurized according to the following conditions:

[0044] After catalyst loading and nitrogen circulation system establishment, the bed temperature of all reactors was raised to 440°C under a nitrogen atmosphere. When the CO+CO2 concentration in the circulating gas exceeded 0.6% by volume, the inlet temperature of each reactor was lowered to 170°C at a rate of 15°C / h, and fresh hydrogen was introduced and the reaction system pressure was gradually increased. Fresh hydrogen was introduced into the hydrotreating reaction zone inlet, replacing gas from the system until the hydrogen gas volume fraction in the circulating hydrogen reached 95%, completing the hydrogen replacement. Then, when the system pressure reached the design pressure of 8 MPa, low-nitrogen start-up oil was introduced. After the oil was introduced, sulfiding agent injection began, raising the reactor inlet temperature to 225°C at a rate of 10°C / h. When the hydrogen sulfide concentration in the circulating hydrogen exceeded 1000 ppm, the reactor inlet temperature was increased to 240°C at a rate of 8°C / h. Once the temperature was reached, constant-temperature sulfidation was initiated for 8 hours, maintaining the hydrogen sulfide concentration between 1000 and 6000 ppm. After constant-temperature sulfidation is complete, ammonia injection passivation begins. The sulfiding agent injection rate is adjusted to maintain a hydrogen sulfide concentration in the circulating hydrogen of 5,000 to 15,000 ppm. The inlet temperature of the two-stage reactor is steadily increased at a rate of 5°C / h. When the ammonia content in the high-density acidic water exceeds 8,000 ppm, it is considered that ammonia has penetrated the catalytic bed. Before ammonia breakthrough, the maximum catalyst bed temperature in the hydrocracking reaction zone is controlled to be ≤ 270°C. After ammonia breakthrough, the passivating agent injection rate is adjusted to maintain the ammonia concentration in the high-density acidic water ≤ 8,000 ppm. The reactor inlet temperature is continuously increased at a rate of 5°C / h to 350°C, and constant-temperature sulfidation is continued for 3 hours.

[0045] After the constant temperature is completed, the inlet temperature of the refining reactor is lowered to 260°C, and the raw oil is prepared to be switched. The evaluation experiment is carried out for 2500 hours according to the conditions in Table 3.

[0046] Example 2

[0047] Hydrorefining Catalyst: Alumina was mixed with water, nitric acid, and graphite (average particle size 20 μm), and then citric acid was added as a binder. The mixture was extruded and dried at 60°C for 10 hours to obtain a shaped support. The cobalt and molybdenum active components were impregnated using ammonium molybdate hexahydrate and cobalt nitrate hexahydrate as the molybdenum source via excess volume co-impregnation. The mixture was dried at 125°C for 8 hours and calcined at 300°C under a nitrogen atmosphere for 3 hours to obtain the finished hydrorefining catalyst F-2. The prepared catalyst had an alumina content of 64%, a graphite content of 7%, a cobalt content calculated as CoO of 4.5% by weight, and a molybdenum content calculated as MoO3 of 24.5% by weight, based on the mass of the catalyst.

[0048] Hydrocracking catalyst: Beta molecular sieve, alumina, water, nitric acid, and graphite (average particle size 20 μm) were mixed, and citric acid was added as a binder. The mixture was extruded and dried at 60°C for 12 hours to obtain a shaped carrier. Molybdenum and nickel active metals were impregnated separately using nickel nitrate hexahydrate as the nickel source and ammonium molybdate hexahydrate as the molybdenum source using an equal volume impregnation method. The mixture was dried at 130°C for 4 hours and calcined at 320°C under a nitrogen atmosphere for 4 hours to obtain the finished hydrocracking catalyst C-2. The catalyst contained 7% beta molecular sieve, 68% alumina, 3% graphite, 5.0% nickel as NiO, and 17.0% molybdenum as MoO3, based on the catalyst mass.

[0049] The hydrotreating catalyst and hydrocracking catalyst prepared above were respectively loaded into the hydrotreating reactor and the hydrocracking reactor, and sulfurized according to the following conditions:

[0050] After catalyst loading and nitrogen recirculation system establishment, the bed temperature of all reactors was raised to 440°C under a nitrogen atmosphere. When the CO+CO2 concentration in the circulating gas exceeded 1.0% by volume, the inlet temperature of each reactor was lowered to 170°C at a rate of 15°C / h, and fresh hydrogen was introduced, with the reaction system pressure gradually increased. Fresh hydrogen was introduced into the hydrotreating reaction zone inlet, displacing the gas from the system until the hydrogen volume fraction in the circulating hydrogen reached 98%, completing the hydrogen replacement. When the system pressure reached the design pressure of 8 MPa, low-nitrogen start-up oil was introduced. After oil introduction, sulfiding agent injection began. The reactor inlet temperature was raised to 220°C at a rate of 10°C / h. When the hydrogen sulfide concentration in the circulating hydrogen exceeded 1000 ppm, the reactor inlet temperature was increased to 245°C at a rate of 8°C / h. Once the temperatures were reached, constant-temperature sulfiding was initiated for 8 hours, maintaining the hydrogen sulfide concentration between 1000 and 6000 ppm. After constant-temperature sulfidation is complete, ammonia injection passivation begins. The sulfiding agent injection rate is adjusted to maintain a hydrogen sulfide concentration in the circulating hydrogen of 5,000 to 15,000 ppm. The inlet temperature of the two-stage reactor is steadily increased at a rate of 5°C / h. When the ammonia content in the high-density acidic water exceeds 8,000 ppm, it is considered that ammonia has penetrated the catalytic bed. Before ammonia breakthrough, the maximum catalyst bed temperature in the hydrocracking reaction zone is controlled to be ≤ 270°C. After ammonia breakthrough, the passivating agent injection rate is adjusted to maintain the ammonia concentration in the high-density acidic water ≤ 8,000 ppm. The reactor inlet temperature is continuously increased at a rate of 5°C / h to 350°C, and constant-temperature sulfidation is continued for 3 hours.

[0051] After the constant temperature is completed, the inlet temperature of the refining reactor is lowered to 270°C, and the feed oil is prepared to be switched. The evaluation experiment is carried out for 2500 hours according to the conditions in Table 3.

[0052] Example 3

[0053] Hydrorefining catalyst: Alumina was mixed with graphite (average particle size 25 μm), water, and nitric acid, and then citric acid was added as a binder. The mixture was extruded and dried at 70°C for 4 hours to obtain a shaped carrier. The cobalt and molybdenum active components were impregnated using ammonium molybdate hexahydrate and cobalt nitrate hexahydrate as the molybdenum source via excess volume co-impregnation. The mixture was dried at 140°C for 4 hours and calcined at 310°C under a nitrogen atmosphere for 4 hours to obtain the finished hydrorefining catalyst F-3. The prepared catalyst contained 70% alumina, 4% graphite, 4.0% cobalt (as CoO), and 22.0% molybdenum (as MoO3), based on the catalyst mass.

[0054] Hydrocracking catalyst: Beta molecular sieve, alumina, water, nitric acid, and graphite (average particle size 25 μm) were mixed, and citric acid was added as a binder. The mixture was extruded and dried at 75°C for 4 hours to obtain a shaped support. Molybdenum and nickel active metals were impregnated separately using nickel nitrate hexahydrate as the nickel source and ammonium molybdate hexahydrate as the molybdenum source using an equal volume impregnation method. The mixture was dried at 130°C for 4 hours and calcined at 310°C under a nitrogen atmosphere for 4 hours to obtain the finished hydrocracking catalyst C-3. The catalyst contained 6% Beta molecular sieve, 65% alumina, 4% graphite, 5.0% nickel (as NiO), and 20.0% molybdenum (as MoO3), based on the catalyst's mass.

[0055] The hydrotreating catalyst and hydrocracking catalyst prepared above were respectively loaded into the hydrotreating reactor and the hydrocracking reactor, and sulfurized according to the following conditions:

[0056] After catalyst loading and nitrogen circulation system establishment, the bed temperature of all reactors was raised to 460°C under a nitrogen atmosphere. When the CO+CO2 concentration in the circulating hydrogen exceeded 1.2% by volume, the inlet temperature of each reactor was lowered to 180°C at a rate of 12°C / h, and fresh hydrogen was introduced and the reaction system pressure was gradually increased. Fresh hydrogen was introduced into the hydrorefining reaction zone inlet to displace gas from the system until the hydrogen gas volume fraction in the circulating hydrogen reached 95%, completing the hydrogen replacement. Then, when the system pressure reached the design pressure of 8 MPa, low-nitrogen start-up oil was introduced. After the oil was introduced, sulfiding agent injection began, raising the reactor inlet temperature to 230°C at a rate of 10°C / h. When the hydrogen sulfide concentration in the circulating hydrogen exceeded 1000 ppm, the reactor inlet temperature was increased to 240°C at a rate of 8°C / h. Once the temperature was reached, constant-temperature sulfidation was initiated for 8 hours, maintaining the hydrogen sulfide concentration between 1000 and 6000 ppm. After constant-temperature sulfiding is completed, ammonia injection passivation begins. The sulfiding agent injection rate is adjusted to maintain a hydrogen sulfide concentration in the circulating hydrogen of 5,000 to 15,000 ppm. The inlet temperature of the two-stage reactor is steadily increased at a rate of 5°C / h. When the ammonia content in the high-density acidic water exceeds 8,000 ppm, it is considered that ammonia has penetrated the catalytic bed. Before ammonia breakthrough, the maximum catalyst bed temperature in the hydrocracking reaction zone is controlled to be ≤ 270°C. After ammonia breakthrough, the passivating agent injection rate is adjusted to maintain the ammonia concentration in the high-density acidic water ≤ 8,000 ppm. The reactor inlet temperature is continuously increased at a rate of 5°C / h to 360°C, and constant-temperature sulfiding is continued for 4 hours.

[0057] After the constant temperature is completed, the inlet temperature of the refining reactor is lowered to 280°C, and the raw oil is prepared to be switched. The evaluation experiment is carried out for 2500 hours according to the conditions in Table 3.

[0058] Example 4

[0059] Hydrorefining catalyst: Alumina was mixed with graphite (average particle size 30 μm), water, and nitric acid, and then citric acid was added as a binder. The mixture was extruded and dried at 60°C for 12 hours to obtain a shaped carrier. The cobalt and molybdenum active components were impregnated using ammonium molybdate hexahydrate and cobalt nitrate hexahydrate as the molybdenum source via excess volume co-impregnation. The mixture was dried at 140°C for 4 hours and calcined at 330°C under a nitrogen atmosphere for 4 hours to obtain the finished hydrorefining catalyst F-4. The prepared catalyst had an alumina content of 63%, a graphite content of 6%, a cobalt content calculated as CoO of 4.0% by weight, and a molybdenum content calculated as MoO3 of 27.0% by weight, based on the catalyst mass.

[0060] Hydrocracking catalyst: Beta molecular sieve, alumina, water, nitric acid, and graphite (average particle size 30 μm) were mixed, and citric acid was added as a binder. The mixture was extruded and dried at 80°C for 4 hours to obtain a shaped support. Molybdenum and nickel active metals were impregnated separately using nickel nitrate hexahydrate as the nickel source and ammonium molybdate hexahydrate as the molybdenum source using an equal volume impregnation method. The mixture was dried at 140°C for 5 hours and calcined at 340°C under a nitrogen atmosphere for 3 hours to obtain the finished hydrocracking catalyst C-4. The catalyst contained 8% Beta molecular sieve, 62% alumina, 5% graphite, 4.0% nickel (as NiO), and 21.0% molybdenum (as MoO3), based on the catalyst mass.

[0061] The hydrorefining catalyst and hydrocracking catalyst prepared above were respectively loaded into the hydrorefining and hydrocracking reactors and sulfided according to the following conditions:

[0062] After catalyst loading and nitrogen circulation system establishment, the bed temperature of all reactors was raised to 480°C under a nitrogen atmosphere. When the CO+CO2 concentration in the circulating gas exceeded 1.5% by volume, the inlet temperature of each reactor was lowered to 160°C at a rate of 15°C / h, and fresh hydrogen was introduced, with the reaction system pressure gradually increased. Fresh hydrogen was introduced from the replacement system at the inlet of the hydrotreating reaction zone until the hydrogen gas volume fraction in the circulating hydrogen reached 95%, completing the hydrogen replacement. When the system pressure reached the design pressure of 8 MPa, low-nitrogen start-up oil was introduced. After the oil was introduced, sulfiding agent injection began. The reactor inlet temperature was raised to 230°C at a rate of 10°C / h. When the hydrogen sulfide concentration in the circulating hydrogen exceeded 1000 ppm, the reactor inlet temperature was increased to 240°C at a rate of 8°C / h. Once the temperatures were reached, constant-temperature sulfiding was initiated for 8 hours, maintaining the hydrogen sulfide concentration between 1000 and 6000 ppm. After constant-temperature sulfidation is complete, ammonia injection passivation begins. The sulfiding agent injection rate is adjusted to maintain a hydrogen sulfide concentration in the circulating hydrogen of 5,000 to 15,000 ppm. The inlet temperature of the two-stage reactor is steadily increased at a rate of 5°C / h. When the ammonia content in the high-density acidic water exceeds 8,000 ppm, it is considered that ammonia has penetrated the catalytic bed. Before ammonia breakthrough, the maximum catalyst bed temperature in the hydrocracking reaction zone is controlled to be ≤ 270°C. After ammonia breakthrough, the passivating agent injection rate is adjusted to maintain the ammonia concentration in the high-density acidic water ≤ 8,000 ppm. The reactor inlet temperature is continuously increased at a rate of 5°C / h to 370°C, and constant-temperature sulfidation is continued for 2 hours.

[0063] After the constant temperature is completed, the inlet temperature of the refining reactor is lowered to 270°C, and the feed oil is prepared to be switched. The evaluation experiment is carried out for 2500 hours according to the conditions in Table 3.

[0064] Example 5

[0065] The preparation methods of the hydrorefining catalyst and the hydrocracking catalyst involved in this embodiment are the same as those in Example 1.

[0066] The hydrotreating catalyst and hydrocracking catalyst prepared above were respectively loaded into the hydrotreating reactor and the hydrocracking reactor, and sulfurized according to the following conditions:

[0067] After catalyst loading and nitrogen circulation system establishment, the bed temperature of all reactors was raised to 440°C under a nitrogen atmosphere. When the CO+CO2 concentration in the circulating gas exceeded 0.6% by volume, the inlet temperature of each reactor was lowered to 170°C at a rate of 15°C / h, and fresh hydrogen was introduced and the reaction system pressure was gradually increased. Fresh hydrogen was introduced into the hydrotreating reaction zone inlet, replacing gas from the system until the hydrogen gas volume fraction in the circulating hydrogen reached 95%, completing the hydrogen replacement. Then, when the system pressure reached the design pressure of 8 MPa, low-nitrogen start-up oil was introduced. After the oil was introduced, sulfiding agent injection began, raising the reactor inlet temperature to 225°C at a rate of 10°C / h. When the hydrogen sulfide concentration in the circulating hydrogen exceeded 1000 ppm, the reactor inlet temperature was increased to 240°C at a rate of 8°C / h. Once the temperature was reached, constant-temperature sulfidation was initiated for 8 hours, maintaining the hydrogen sulfide concentration between 1000 and 6000 ppm. After constant-temperature sulfiding is complete, ammonia injection passivation begins. The sulfiding agent injection rate is adjusted to maintain a hydrogen sulfide concentration in the circulating hydrogen of 5,000 to 15,000 ppm. The inlet temperature of the two-stage reactor is steadily increased at a rate of 5°C / h. When the ammonia content in the high-density acidic water exceeds 8,000 ppm, it is considered that ammonia has penetrated the catalytic bed. Before ammonia breakthrough, the maximum catalyst bed temperature in the hydrocracking reaction zone is controlled to be ≤ 270°C. After ammonia breakthrough, the passivating agent injection rate is adjusted to maintain the ammonia concentration in the high-density acidic water ≤ 8,000 ppm. The reactor inlet temperature is continuously increased at a rate of 5°C / h to 340°C, and constant-temperature sulfiding is continued for 3 hours.

[0068] After the constant temperature is completed, the inlet temperature of the refining reactor is lowered to 260°C, and the raw oil is prepared to be switched. The evaluation experiment is carried out for 2500 hours according to the conditions in Table 3.

[0069] Comparative Example 1

[0070] The hydrotreating reactor of the coal-to-liquid hydrocracking unit is loaded with industrial FF-66 catalyst, while the hydrocracking reactor is loaded with industrial FC-14 hydrocracking catalyst. The sulfurization is carried out according to the following conventional sulfurization method:

[0071] After the catalyst loading and the establishment of the circulating hydrogen system in the coal-to-liquid unit were completed, the reaction inlet temperature was increased to 170°C and the high-pressure was controlled to the design pressure of 8MPa.

[0072] A sulfiding agent is injected into the reaction system via a high-pressure sulfur injection pump at a rate of 0.5 t / h. After hydrogen sulfide penetrates the bed, the reactor inlet temperature is raised to 230°C at a rate of 8°C / h. Sulfidation is carried out at a constant temperature for 8 hours, during which the hydrogen sulfide concentration in the circulating hydrogen should be maintained between 1000 and 5000 ppm. The reactor inlet temperature is then steadily raised to 310°C at a rate of 6°C / h. Sulfidation is carried out at a constant temperature for 4 hours. During this period, the hydrogen sulfide concentration in the circulating hydrogen should be maintained between 6000 and 10000 ppm.

[0073] After the constant temperature at 310°C was completed, the reactor inlet temperature was lowered to 260°C, the feedstock oil was introduced, the reaction system operating conditions were adjusted as shown in Table 3, and a 2500-h stability test was carried out.

[0074] Comparative Example 2

[0075] The hydrotreating reactor of the coal-to-liquid hydrocracking unit is loaded with industrial FF-66 catalyst, while the hydrocracking reactor is loaded with industrial FC-76 hydrocracking catalyst. The sulfurization is carried out according to the following conventional sulfurization method:

[0076] After the catalyst loading and circulating hydrogen system establishment work was completed in the coal-to-liquid unit, the reaction inlet temperature was increased to 170°C and the high-pressure was controlled to the design pressure of 8MPa.

[0077] A sulfiding agent is injected into the reaction system via a high-pressure sulfur injection pump at a rate of 0.5 t / h. After hydrogen sulfide penetrates the bed, the reactor inlet temperature is raised to 230°C at a rate of 8°C / h. Sulfidation is carried out at a constant temperature for 8 hours, during which the hydrogen sulfide concentration in the circulating hydrogen should be maintained between 1000 and 5000 ppm. The reactor inlet temperature is then steadily raised to 310°C at a rate of 6°C / h. Sulfidation is carried out at a constant temperature for 4 hours. During this period, the hydrogen sulfide concentration in the circulating hydrogen should be maintained between 6000 and 10000 ppm.

[0078] After the constant temperature at 310°C was completed, the reactor inlet temperature was lowered to 260°C, the feedstock oil was introduced, the reaction system operating conditions were adjusted as shown in Table 3, and a 2500-h stability test was carried out.

[0079] Comparative Example 3

[0080] The preparation methods of the hydrorefining catalyst and the hydrocracking catalyst involved in this comparative example are as follows:

[0081] Hydrorefining Catalyst: Alumina was mixed with water, nitric acid, and a binder of citric acid and extruded into strips. The strips were then dried at 50°C for 10 hours to obtain a shaped carrier. The cobalt and molybdenum active components were impregnated using ammonium molybdate hexahydrate and cobalt nitrate hexahydrate as the molybdenum source via excess volume co-impregnation. The mixture was dried at 120°C for 6 hours and calcined at 500°C in air for 4 hours to obtain the finished hydrorefining catalyst F-5. The prepared catalyst had an alumina content of 64%, a cobalt content calculated as CoO of 5.0% by weight, and a molybdenum content calculated as MoO3 of 31.0% by weight, based on the mass of the catalyst.

[0082] Hydrocracking catalyst: Beta molecular sieve, alumina, water, and nitric acid were mixed, and citric acid was added as a binder. After extrusion, the mixture was dried at 55°C for 12 hours to obtain a shaped support. Nickel nitrate hexahydrate was used as the nickel source, and ammonium molybdate hexahydrate was used as the molybdenum source. Molybdenum and nickel active metals were impregnated separately using an equal volume impregnation method. The mixture was dried at 120°C for 10 hours and calcined at 500°C in an air atmosphere for 5 hours to obtain the finished hydrocracking catalyst C-5. The catalyst contained 6% beta molecular sieve, 69% alumina, 5.0% nickel as NiO, and 20.0% molybdenum as MoO3, based on the catalyst mass.

[0083] After catalyst loading and nitrogen circulation system establishment, the bed temperature of all reactors was raised to 440°C under a nitrogen atmosphere. When the CO+CO2 concentration in the circulating gas exceeded 0.6% by volume, the inlet temperature of each reactor was lowered to 170°C at a rate of 15°C / h, and fresh hydrogen was introduced and the reaction system pressure was gradually increased. Fresh hydrogen was introduced into the hydrotreating reaction zone inlet, replacing gas from the system until the hydrogen gas volume fraction in the circulating hydrogen reached 95%, completing the hydrogen replacement. Then, when the system pressure reached the design pressure of 8 MPa, low-nitrogen start-up oil was introduced. After the oil was introduced, sulfiding agent injection began, raising the reactor inlet temperature to 225°C at a rate of 10°C / h. When the hydrogen sulfide concentration in the circulating hydrogen exceeded 1000 ppm, the reactor inlet temperature was increased to 240°C at a rate of 8°C / h. Once the temperature was reached, constant-temperature sulfidation was initiated for 8 hours, maintaining the hydrogen sulfide concentration between 1000 and 6000 ppm. After constant-temperature sulfidation is complete, ammonia injection passivation begins. The sulfiding agent injection rate is adjusted to maintain a hydrogen sulfide concentration in the circulating hydrogen of 5,000 to 15,000 ppm. The inlet temperature of the two-stage reactor is steadily increased at a rate of 5°C / h. When the ammonia content in the high-density acidic water exceeds 8,000 ppm, it is considered that ammonia has penetrated the catalytic bed. Before ammonia breakthrough, the maximum catalyst bed temperature in the hydrocracking reaction zone is controlled to be ≤ 270°C. After ammonia breakthrough, the passivating agent injection rate is adjusted to maintain the ammonia concentration in the high-density acidic water ≤ 8,000 ppm. The reactor inlet temperature is continuously increased at a rate of 5°C / h to 350°C, and constant-temperature sulfidation is continued for 3 hours.

[0084] After the constant temperature is completed, the inlet temperature of the refining reactor is lowered to 260°C, and the raw oil is prepared to be switched. The evaluation experiment is carried out for 2500 hours according to the conditions in Table 3.

[0085] Table 4 Example test results

[0086] Liquid product yield after 2500h, wt% Example 1 Example 2 Example 3 Example 4 Example 5 naphtha 29.5 30.4 30.5 29.8 28.8 diesel fuel 41.8 41.2 40.0 41.9 42.5 Naphtha bromine index, Br mg / 100g 56 70 75 51 90 Hydrogen consumption, wt% 0.50 0.51 0.52 0.49 0.53 Deactivation rate of hydrotreating catalyst, ℃ / d 0.03 0.03 0.03 0.02 0.04 Deactivation rate of hydrocracking catalyst, ℃ / d 0.02 0.03 0.04 0.02 0.05

[0087] Table 5 Comparative Example Test Results

[0088] Liquid product yield after 2500h, wt% Comparative Example 1 Comparative Example 2 Comparative Example 3 naphtha 26.9 25.4 26.0 diesel fuel 38.2 37.1 38.0 Naphtha bromine index, Br mg / 100g 200 215 205 Hydrogen consumption, wt% 0.71 0.76 0.75 Deactivation rate of hydrotreating catalyst, ℃ / d 0.07 0.08 0.07 Deactivation rate of hydrocracking catalyst, ℃ / d 0.07 0.07 0.07

[0089] It can be seen from the experimental results of the comparative examples and the embodiments that, under the condition of controlling the same temperature rise, when the coal-to-liquid hydrocracking process of the present invention is used, not only the diesel and naphtha yields produced are high, but also the naphtha bromine index and hydrogen consumption are low, and the catalyst deactivation rate is slow.

Claims

1. A coal-to-liquid hydrocracking method, wherein: The method comprises a hydrorefining reaction zone and a hydrocracking reaction zone, wherein the hydrorefining reaction zone is filled with a hydrorefining catalyst containing graphite, and the graphite content is 2% to 10% based on the mass of the hydrorefining catalyst containing graphite; and the hydrocracking reaction zone is filled with a hydrocracking catalyst containing graphite, and the graphite content is 2% to 10% based on the mass of the hydrocracking catalyst containing graphite. The start-up process comprises: after the catalyst loading is completed and the inert gas circulation system is established, the temperature of each catalyst bed is increased to 440 to 500° C. under an inert gas environment; when the total concentration of CO and CO2 in the circulating gas is detected to reach 0.6% to 2.0% by volume, the inlet temperature of each reactor is reduced to 150 to 190° C. at a rate of ≯20° C. / h; then, the device is subjected to hydrogen replacement and sulfurization; and after the sulfurization is completed, the feed is switched to introduce crude oil and normal production is resumed; The hydrocracking catalyst containing graphite comprises an active metal component and a carrier, wherein the active metal component is a Group VIII and Group VIB metal, and the carrier comprises Beta molecular sieve, alumina and graphite; The sulfidation adopts wet sulfidation, and the wet sulfidation includes a first constant temperature sulfidation and a second constant temperature sulfidation. The first constant temperature sulfidation is to increase the inlet temperature of the hydrofining reaction zone and the hydrocracking reaction zone to a constant temperature of 240-260°C at a rate of ≯10°C / h, and then to perform constant temperature sulfidation for 6-10 hours at the constant temperature. During this sulfidation stage, the mass concentration of hydrogen sulfide in the circulating hydrogen is controlled to be maintained at 1000-6000ppm. The second constant temperature sulfidation is to increase the inlet temperature of the hydrofining reaction zone and the hydrocracking reaction zone to a second constant temperature of 340-380°C at a rate of ≯6°C / h, and to perform sulfidation at the second constant temperature for 2-4 hours. During this sulfidation stage, the mass concentration of hydrogen sulfide in the circulating hydrogen is controlled to be maintained at 5000-15000ppm.

2. The method according to claim 1, characterized in that Based on the mass of the hydrotreating catalyst containing graphite, the graphite content is 2% to 6%.

3. The method according to claim 1, characterized in that Based on the mass of the hydrocracking catalyst containing graphite, the content of graphite is 2% to 6%.

4. The method according to claim 1, wherein When the total concentration of CO and CO2 in the circulating gas is detected to reach 0.6v%~2.0v%, the inlet temperature of each reactor is reduced to 150~190℃ at a rate of 5~20℃ / h.

5. The method according to claim 1, wherein The graphite-containing hydrorefining catalyst comprises a carrier and an active metal. The carrier comprises alumina and graphite, and the active metal is a Group VIII or Group VIB metal.

6. The method according to claim 5, characterized in that In the graphite-containing hydrorefining catalyst, the Group VIII metal is Co and / or Ni, and the Group VIB metal is W and / or Mo.

7. The method according to claim 5, characterized in that Based on the weight of the hydrorefining catalyst, the content of alumina is 15% to 65%, the content of graphite is 2% to 10%, the content of binder is less than 8%, and the content of active metals calculated as oxides is 12% to 50%, of which the content of Group VIII metals calculated as oxides is 1.5% to 10%, and the content of Group VIB metals calculated as oxides is 10% to 40%.

8. The method according to claim 1, characterized in that In the hydrocracking catalyst containing graphite, the Group VIII metal is Co and / or Ni, and the Group VIB metal is W and / or Mo.

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

10. The method according to claim 1, characterized in that Based on the mass of the hydrocracking catalyst containing graphite, the molecular sieve content is 5% to 10%, the alumina content is 20% to 60%, the graphite content is 2% to 10%, and the binder content is less than 10%.

11. The method according to claim 1 or 8, characterized in that Based on the mass of the hydrocracking catalyst containing graphite, the content of the Group VIII metal calculated as oxide is 1% to 8%, and the content of the Group VIB metal calculated as oxide is 14% to 32%.

12. The method according to claim 1, characterized in that The hydrogen replacement is completed when the hydrogen gas volume fraction in the circulating hydrogen is adjusted to 92% to 99%.

13. The method according to claim 1, wherein During the hydrogen replacement, the inlet temperatures of the hydrofining reaction zone and the hydrocracking reaction zone are first lowered to 140-190°C.

14. The method according to claim 1, wherein The first constant temperature sulfurization increases the inlet temperature of the hydrofining reaction zone and the hydrocracking reaction zone to a constant temperature at a rate of 2-10°C / h.

15. The method according to claim 1, characterized in that The second constant temperature sulfurization increases the inlet temperature of the hydrofining reaction zone and the hydrocracking reaction zone to a second constant temperature at a rate of 2-6°C / h.

16. The method according to claim 14, characterized in that The wet sulfiding process allows hydrogen sulfide to penetrate the catalyst bed before entering the sulfiding stage. The process is as follows: low-nitrogen oil is introduced from the inlet of the hydrorefining reaction zone to wet the catalyst, a closed-loop circulation of the low-nitrogen oil is established, and then a sulfiding agent is gradually injected into the low-nitrogen oil. The temperature is gradually increased and the catalyst bed temperature in the hydrorefining reaction zone and the hydrocracking reaction zone is controlled to not exceed 230° C. until hydrogen sulfide penetrates the entire catalyst bed. The low-nitrogen oil has a nitrogen content of 100 to 800 ppm by mass.

17. The method according to claim 16, characterized in that The conditions for hydrogen sulfide to penetrate the entire catalyst bed are as follows: the mass concentration of hydrogen sulfide in the circulating hydrogen is above 1000 ppm.

18. The method according to claim 17, characterized in that The conditions for hydrogen sulfide to penetrate the entire catalyst bed are as follows: the mass concentration of hydrogen sulfide in the circulating hydrogen is 1000~10000ppm.

19. The method according to claim 14, wherein After the first constant temperature sulfiding is completed, a passivating agent is injected into the inlet of the hydrocracking reaction zone, and the injection rate of the sulfiding agent is adjusted so that the hydrogen sulfide concentration in the circulating hydrogen reaches 5000~15000ppm. The ammonia content in the acidic water discharged from the bottom of the high-pressure separator is monitored. Before ammonia breakthrough, the maximum temperature of the catalyst bed in the hydrocracking reaction zone is controlled to be ≯270℃. After ammonia breakthrough, the passivating agent injection rate is adjusted to maintain the ammonium ion mass concentration in the acidic water discharged from the bottom of the high-pressure separator at 1000~8000ppm. The addition of the passivating agent is stopped until the sulfiding process is completed.

20. The method according to claim 19, characterized in that The condition for ammonia breakthrough is that the mass concentration of ammonium ions in the acidic water discharged from the bottom of the high-pressure separator reaches 6000 ppm or more.

21. The method according to claim 19, wherein The condition for ammonia breakthrough is that the mass concentration of ammonium ions in the acidic water discharged from the bottom of the high-pressure separator reaches 8000-15000 ppm.

22. The method according to claim 19, wherein The vulcanizing agent is at least one of dimethyl disulfide and carbon disulfide.

23. The method according to claim 19, wherein The passivating agent is selected from anhydrous liquid ammonia and / or an organic passivating agent, and the organic passivating agent is one or more of cyclohexylamine, aniline, and n-butylamine.

24. The method according to claim 16, wherein The low nitrogen oil is atmospheric distillate oil with a density of 0.76-0.86 g / cm 3 The initial distillation point is 100~250℃, and the final distillation point is 280~408℃.

25. The method according to claim 1, wherein The coal-to-liquid hydrocracking method comprises: coal-to-liquid feedstock is mixed with hydrogen and then enters a hydrorefining reaction zone to undergo deolefination and deoxygenation reactions; the effluent from the hydrorefining reaction enters a first-stage separation system to obtain light components and heavy components; the heavy components enter a hydrocracking reaction zone to undergo an isomerization chain scission reaction; and the effluent from the hydrocracking reaction enters a second-stage separation system to obtain light naphtha, heavy naphtha, diesel, and wax oil products.

26. The method according to claim 25, characterized in that The cut point between the light components and heavy components separated by the first separation system is 165~210℃.

27. The method according to claim 1, wherein The coal-to-liquid has an initial boiling point of 20-100° C., a final boiling point of 300-400° C., a sulfur content of ≯50 ppm by mass, and a nitrogen content of ≯20 ppm by mass.

28. The method according to claim 1, wherein During normal production, the reaction conditions in the hydrotreating reaction zone are as follows: reaction pressure of 6.0-15.0 MPa, average reaction temperature of 270-340°C, hydrogen-to-oil volume ratio of 400:1-2200:1, liquid hourly space velocity of 0.5-4.0 h -1 The reaction conditions of the hydrocracking reaction zone are: reaction pressure of 6.0-15.0 MPa; average reaction temperature of 280-400°C; hydrogen-oil volume ratio of 400:1-2200:1, liquid hourly space velocity of 0.5-5.0 h -1 .

29. The method according to claim 28, characterized in that During normal production, the reaction conditions of the hydrorefining reaction zone are as follows: reaction pressure of 7.0-12.0 MPa, average reaction temperature of 280-320°C, hydrogen-to-oil volume ratio of 600:1-1400:1; the reaction conditions of the hydrocracking reaction zone are as follows: reaction pressure of 7.0-12.0 MPa, average reaction temperature of 300-350°C, hydrogen-to-oil volume ratio of 800:1-1600:1, liquid hourly volume space velocity of 1.0-4.0 h -1 .

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