A method for producing gasoline by catalytic hydrogenation of diesel

Abstract

The present invention discloses a method for producing gasoline by catalytic diesel hydrogenation. The method comprises: (1) passivated raw oil, a passivation aid and hydrogen enter a hydrogenation reactor, and sequentially contact with a hydrogenation pretreatment catalyst and a hydrocracking catalyst to perform a passivation reaction; (2) after the passivation reaction is completed, the passivated raw oil is switched to catalytic diesel in batches until the final feed is a mixture of catalytic diesel and the passivation aid; (3) the mixture of catalytic diesel and the passivation aid and hydrogen sequentially contact with the passivated hydrogenation pretreatment catalyst and the hydrocracking catalyst, and perform a hydrogenation pretreatment reaction and a hydrocracking reaction, respectively, to obtain gasoline. The present invention can shorten the adjustment time for the octane number, sulfur content, etc. of the gasoline product to reach a qualified level in the initial stage of start-up, and can reduce investment.

Description

Technical Field

[0001] The invention belongs to the field of petroleum refining, and relates to a method for producing gasoline by catalytic diesel hydrogenation, and specifically to a hydrogenation conversion process for rapidly producing high-octane gasoline by using catalytic diesel. Background Art

[0002] Catalytic cracking (FCC) technology is one of the main technological means for lightening heavy oil, and it occupies a relatively important position in oil refining enterprises in various countries in the world. In recent years, as the raw materials processed by catalytic cracking units have become heavier and inferior, the quality of catalytic cracking products, especially catalytic cracking light cycle oil (LCO), has deteriorated. Catalytic cracking light cycle oil has high sulfur and aromatic content and poor engine ignition performance. It is mainly used for blending fuel oil, non-automotive diesel and heating oil abroad. Converting aromatics into high-octane fuel is a method of moderate cracking and shallow hydrogenation. The depth of hydrogenation during catalytic diesel hydrogenation conversion determines the loss of octane number.

[0003] At present, there are reports on the use of hydrocracking technology to convert catalytic cracking light cycle oil into ultra-low sulfur diesel and high-octane gasoline blending components. For example, at the 1995 NPRA Annual Meeting, DaVId A. Pappal et al. introduced the MAK-LCO technology developed by Mobil, Akzo Nobel and MW Kellogg; at the 2005 NPRA Annual Meeting, Vasant P. Thakkar et al. introduced the LCO UnicrackingTM technology and HC-190 special catalyst developed by UOP. It is reported that both of the above technologies can convert low-value catalytic cycle oil components into high-octane gasoline components and high-quality diesel blending components. In addition, at the 2007 NPRA Annual Meeting, UOP introduced its new LCO-X technology, which achieved a new way to increase the production of aromatics using catalytic light cycle oil through the hydroconversion-selective alkyl transfer route. It can be seen that the utilization of relatively low-value catalytic light cycle oil has been upgraded from the initial modification to the level of producing higher value-added oil and chemical products.

[0004] At present, the catalytic diesel hydroconversion units in operation in China all have the problem that the initial activity of the catalyst is high at the beginning of the unit operation, the hydrodesulfurization, denitrogenation and dearomatization capabilities are strong, the hydrocracking catalyst has high hydrogenation performance, and the initial matching of hydrogenation activity and cracking activity is poor. This shows that the octane number of the gasoline component is not high at the beginning of the operation, which cannot meet the production requirements. It needs to be stabilized for a long time before the octane number of the produced gasoline meets the requirement of not less than 90.

[0005] US10162310 discloses a two-phase hydrogenation pretreatment technology. The process is that fresh raw oil, circulating oil and hydrogen pass through a hydrogen mixing device to dissolve hydrogen in the oil, and the oil with dissolved hydrogen enters a smaller reactor to contact with the catalyst for hydrogenation reaction to remove impurities in the oil. After the reaction, part of the logistics is circulated to the hydrogen mixing device, and part is discharged from the device as a product. This method uses the raw material and circulating oil to dissolve the required hydrogen in the oil before entering the reactor, and the circulating hydrogen system can be omitted. When this method treats the secondary processed intermediate distillate oil, the impurity removal rate is difficult to meet the standard. In the existing two-phase hydrogenation technology, hydrogen sulfide dissolved in the oil phase is difficult to remove effectively, and it accumulates in the reaction system, causing a serious inhibitory effect on the hydrogenation reaction. Although it can be removed by methods such as stripping, due to being in a high temperature and high pressure system, it will also increase equipment investment and operating costs.

[0006] CN104611050A discloses a catalytic cracking diesel conversion method. The method adopts a single-stage series process, and the hydrofining reaction effluent directly enters the hydrocracking reactor and contacts and reacts with the graded catalyst bed in the cracking reactor; according to the flow direction of the reaction material, the hydrogenation activity of the hydrocracking catalyst tends to decrease. The inventive method can reduce the over-hydrogenation and secondary cracking of cracked naphtha while ensuring the diesel hydrocracking effect, reduce chemical hydrogen consumption, and thus improve the octane number and liquid yield of naphtha. However, the problem of high catalyst activity in the initial stage of the device is not considered, and the device adjustment time cannot be shortened. Summary of the invention

[0007] In view of the shortcomings of the prior art, the purpose of the present invention is to provide a method for producing gasoline by catalytic diesel hydrogenation, especially by directly adopting the existing trickle bed hydrogenation reaction process equipment, the adjustment time for the octane number, sulfur content, etc. of the gasoline product to reach the qualified level in the initial stage of start-up can be shortened, and investment can be reduced.

[0008] The present invention provides a method for producing gasoline by catalytic hydrogenation of diesel, comprising:

[0009] (1) The passivated feedstock oil, the passivation aid and the hydrogen enter the hydrogenation reactor and sequentially contact with the hydrogenation pretreatment catalyst and the hydrocracking catalyst to perform a passivation reaction;

[0010] (2) After the passivation reaction is completed, the passivation feedstock oil is switched to catalytic diesel in batches until the final feedstock is a mixture of catalytic diesel and passivation aid;

[0011] (3) The mixture of catalytic diesel and the passivation aid and hydrogen are contacted with the passivated hydrogenation pretreatment catalyst and the hydrocracking catalyst in sequence to carry out hydrogenation pretreatment reaction and hydrocracking reaction respectively to obtain gasoline.

[0012] Furthermore, the hydrogenation reactor is a trickle bed reactor.

[0013] Furthermore, the hydrotreating catalyst and the hydrocracking catalyst have been subjected to conventional sulfurization or activation treatment before the passivation reaction.

[0014] Furthermore, the initial distillation point of the passivated feedstock oil is 150°C to 350°C, preferably 150°C to 250°C, the final distillation point is 370°C to 420°C, the aromatic content is <30wt%, preferably 15wt% to 25wt%, and the nitrogen content is <200μg / g, preferably 50 to 100μg / g.

[0015] Furthermore, the passivated feedstock oil is a heavy diesel fraction, such as one or more of normal second-line, normal third-line, normal fourth-line, reduced top oil, reduced first-line, straight-run diesel, etc., preferably one or more of normal fourth-line, reduced top oil and reduced first-line, etc.

[0016] Furthermore, the passivation aid is selected from one or more of sulfides such as di-tert-butyl sulfide, di-tert-butyl disulfide, di-tert-butyl polysulfide, tert-nonyl polysulfide, tert-dodecyl polysulfide, diphenyl sulfide, diphenyl disulfide, diphenyl polysulfide, and di-tert-nonyl sulfide.

[0017] Furthermore, the conditions of the passivation reaction are: reaction temperature is 310°C to 380°C, preferably 320°C to 350°C, volume space velocity is 0.2h -1 ~6.0h -1 , preferably 0.6h -1 ~5h -1 The volume ratio of hydrogen to oil is 200:1 to 2000:1, preferably 500:1 to 1500:1, and the total passivation time is 40 to 100 hours.

[0018] Furthermore, the passivation reaction includes two stages. In the first stage, the passivation aid accounts for 0% to 15% of the passivation raw oil mass, preferably 1% to 15%, and the operating time is 16 to 72 hours, preferably 24 to 48 hours; in the second stage, the passivation aid accounts for 16% to 30% of the passivation raw oil mass, preferably 20% to 30%, and the operating time is 12 to 48 hours, preferably 16 to 32 hours.

[0019] Further, the hydrogenation pretreatment catalyst includes a carrier and a hydrogenation active component, wherein the carrier is generally a non-acidic or weakly acidic refractory porous oxide, such as alumina, silicon oxide, amorphous silicon aluminum, titanium oxide, and at least one of a composite oxide or mixed oxide carrier of several elements. The hydrogenation active component is W-Ni, Mo-Ni or W-Mo-Ni. Based on the weight of the hydrogenation pretreatment catalyst, the content of the hydrogenation active component in terms of oxide is generally 15wt% to 50wt%, preferably 18wt% to 45wt%. Among them, the content of nickel oxide is further preferably 2wt% to 8wt%, more preferably 2.6wt% to 6wt%. The hydrogenation pretreatment catalyst can select a suitable commercial catalyst according to the needs of the process flow, can also be prepared according to the existing method, or can be a regenerated catalyst after the deactivated catalyst is regenerated. For example, it can be FHUDS-8 produced by Fushun Branch of Sinopec Catalyst Company.

[0020] Furthermore, the hydrocracking catalyst comprises a carrier and a hydrogenation active component, wherein the carrier is an acidic material, such as Y-type and / or USY molecular sieve. Based on the weight of the hydrocracking catalyst, the content of the carrier is 20wt% to 80wt%, preferably 30wt% to 70wt%; the hydrogenation active component is one or more of W, Mo, Ni and Co, and based on the weight of the hydrocracking catalyst, the content of the hydrogenation active component in terms of oxide is generally 3wt% to 50wt%. The hydrocracking catalyst can be selected from suitable commercial catalysts according to the needs of the process flow, or can be prepared according to existing methods, or can be a regenerated catalyst after a deactivated catalyst is regenerated. For example, it can be FC-24 produced by Fushun Branch of Sinopec Catalyst Company.

[0021] Furthermore, the volume ratio of the hydropretreatment catalyst to the hydrocracking catalyst is 3:7 to 7:3, preferably 4:6 to 6:4.

[0022] Furthermore, the initial distillation point of the catalytic diesel is 65°C to 290°C, preferably 100°C to 200°C, the final distillation point is 330°C to 450°C, the aromatic content of the catalytic diesel fraction is ≥50wt%, preferably 50wt% to 75wt%, and the polycyclic aromatic hydrocarbon content is ≥35wt%, preferably 40wt% to 50wt%.

[0023] Furthermore, the catalytic diesel can be a product of various catalytic cracking processes obtained by processing any base oil, such as one or more of DCC diesel, MIP diesel, RTC diesel, LTAG diesel, FDFCC diesel, CPP diesel and ARGG diesel. The catalytic diesel can also be one or more of mixed aromatics extraction unit raffinate, coal liquefaction diesel, coal tar diesel and shale oil diesel fraction.

[0024] Furthermore, the passivated stock oil is switched to catalytic diesel in batches, specifically, based on the total weight of the passivated stock oil and the catalytic diesel, the proportion of the catalytic diesel is increased by 10% to 35% each time.

[0025] Furthermore, based on the total weight of the passivated raw oil and the catalytic diesel, each time the proportion of the catalytic diesel is increased, the proportion of the passivation aid is reduced by 1% to 5%.

[0026] Furthermore, based on the total weight of the catalytic diesel in the final feed, the content of the passivation aid in the final feed is 5% to 20%.

[0027] Furthermore, the interval between each switching of catalytic diesel is 5 to 15 hours.

[0028] Furthermore, the conditions of the hydrofining reaction include: reaction temperature of 260°C to 430°C, preferably 300°C to 380°C; volume space velocity of 0.4h -1 ~10.0h -1 , preferably 0.6h -1 ~5h -1 ; The volume ratio of hydrogen to oil is 200:1 to 2000:1, preferably 500:1 to 1500:1, and the total pressure is 2.5 to 20.0 MPa, preferably 5.0 to 15.0 MPa.

[0029] Furthermore, the conditions of the hydrocracking reaction include: a reaction temperature of 350°C to 450°C, preferably 355°C to 420°C; a volume space velocity of 0.4h -1 ~16.0h -1 , preferably 0.6h -1 ~6h -1 The volume ratio of hydrogen to oil is 200:1 to 2000:1, preferably 500:1 to 1500:1, and the total pressure is 2.5 to 20.0 MPa, preferably 5.0 to 15.0 MPa.

[0030] Furthermore, the reaction products obtained through the hydrogenation pretreatment reaction and the hydrocracking reaction are separated into gas and liquid, and then enter the fractionation zone to obtain gasoline through stripping and fractionation.

[0031] Furthermore, the sulfur content of the obtained gasoline product is ≯10 μg / g, and the research octane number is ≮92.

[0032] Compared with the prior art, the method of the present invention has the following advantages:

[0033] 1. The present invention adopts the existing catalytic diesel hydroconversion process, strengthens the passivation stage of the catalyst at the initial stage of operation, and adds a passivation aid to achieve the regulation of reducing the activity of the catalyst, thereby achieving a controllable carbon deposition process of the catalyst and shortening the time for the octane number of the gasoline product to reach the qualified level.

[0034] 2. The catalyst passivation process begins with the injection of the passivation aid, which decomposes into small molecular hydrocarbons and H in the hydrogen atmosphere. 2 S, to achieve the effect of changing the total pressure of the reaction system at different periods, realize the step-by-step carbon deposition of the catalyst, avoid excessive carbon deposition of the catalyst, achieve catalyst hydrogenation performance matching, and at the same time avoid the situation where the operating pressure of the circulating hydrogen compressor is too low and the compressor cannot operate.

[0035] 3. The method of the present invention uses a passivation agent to decompose small hydrocarbon molecules and H 2 S, through oil separation and circulating hydrogen desulfurization, the operation can be controlled, excessive enrichment of circulating hydrogen can be avoided, and the operation of the device and the quality of other products will not be affected during normal production.

[0036] 4. In the method of the present invention, the passivation aid is more conducive to the hydrocracking catalyst to exert cracking activity and reduce its hydrogenation performance under the atmosphere of hydrogen.

[0037] 5. The method of the present invention follows the existing process technology, strengthens the initial passivation process, has little process operation change, fully matches the activity, and shortens the time for the octane number of the gasoline product to reach the octane number requirement of the 92 national standard gasoline. DETAILED DESCRIPTION

[0038] The following non-limiting embodiments may enable those skilled in the art to more fully understand the present invention, but are not intended to limit the present invention in any way.

[0039] The hydroconversion pretreatment catalyst used in the embodiments and comparative examples of the present invention is FHUDS-8 developed by Fushun Petrochemical Research Institute and produced by Fushun Branch of Sinopec Catalyst Company; the hydroconversion catalyst is FC-24 developed by Fushun Petrochemical Research Institute and produced by Fushun Branch of Sinopec Catalyst Company. The FHUDS-8 catalyst uses alumina as a carrier and Mo-Ni as an active component; the FC-24 catalyst uses a modified Y-type molecular sieve as a conversion component and W-Ni as an active metal component.

[0040] The conventional one-stage series hydrocracking process used in the embodiments and comparative examples of the present invention has a trickle bed reactor as the reactor.

[0041] Example 1

[0042] The sulfurized FHUDS-8 and FC-24 catalysts were loaded into the hydrogenation reactor at a volume ratio of 5:6. The first stage of catalyst passivation was carried out, and the passivated feedstock oil, passivation aid (TBPS-454, di-tert-butyl polysulfide) and hydrogen entered the hydrogenation reactor. The passivation aid accounted for 13% of the passivated feedstock oil mass. The reaction temperature was adjusted to 345°C in the hydrogenation pretreatment zone and 355°C in the hydrocracking zone. The volume space velocity was controlled to 0.55h -1 After the constant temperature operation, the catalyst was deactivated in the second stage. The amount of deactivator was adjusted to 24% of the mass of the deactivator raw oil. Other conditions remained unchanged and the constant temperature was continued for 26 hours.

[0043] After the passivation constant temperature is completed, based on the total weight of the passivated raw oil and catalytic diesel, switch to 25% catalytic diesel, and reduce the amount of passivation to 21%, and keep the constant temperature for 12 hours; after the constant temperature is completed, increase the proportion of catalytic diesel to 50% again, and reduce the content of passivation agent to 18%, and continue the constant temperature for 12 hours; after the constant temperature is completed, increase the proportion of catalytic diesel to 75% again, and reduce the content of passivation agent to 15%, and continue the constant temperature for 12 hours; after the constant temperature is completed, increase the proportion of catalytic diesel to 100%, and reduce the content of passivation agent to 12%, and continue the constant temperature for 12 hours. After this process is completed, stop adding passivation agent and carry out subsequent hydrofining and hydrocracking reactions.

[0044] Example 2

[0045] The sulfurized FHUDS-8 and FC-24 catalysts were loaded into the hydrogenation reactor at a volume ratio of 5:6. The first stage of catalyst passivation was carried out, and the passivated feedstock oil, passivation aid (TBPS-454, di-tert-butyl polysulfide) and hydrogen entered the hydrogenation reactor. The passivation aid accounted for 15% of the passivated feedstock oil mass. The reaction temperature was adjusted to 345°C in the hydrogenation pretreatment zone and 355°C in the hydrocracking zone. The volume space velocity was controlled to 0.55h -1 After the constant temperature operation, the second stage of catalyst passivation is started. The amount of passivation agent is adjusted to 30% of the passivation raw oil mass. Other conditions remain unchanged and the constant temperature is continued for 32 hours.

[0046] After the passivation constant temperature is completed, based on the total weight of the passivated raw oil and catalytic diesel, switch to 20% catalytic diesel, and reduce the passivator content to 28%, and keep the constant temperature for 12 hours; after the constant temperature is completed, increase the catalytic diesel ratio to 40% again, and reduce the passivator content to 26%, and continue the constant temperature for 12 hours; after the constant temperature is completed, increase the catalytic diesel ratio to 60% again, and reduce the passivator content to 24%, and continue the constant temperature for 12 hours; after the constant temperature is completed, increase the catalytic diesel ratio to 80%, and reduce the passivator content to 22%, and continue the constant temperature for 12 hours; after the constant temperature is completed, increase the catalytic diesel ratio to 100%, and reduce the passivator content to 20%, and continue the constant temperature for 12 hours. After this process is completed, stop adding the passivator and carry out subsequent hydrofining and hydrocracking reactions.

[0047] Example 3

[0048] Compared with Example 1, the only difference is that the passivation aid is changed to tert-nonyl polysulfide.

[0049] Example 4

[0050] Compared with Example 2, the only difference is the conditions of hydrofining and hydrocracking, as shown in Table 2.

[0051] Comparative Example 1

[0052] Compared with Example 1, the only difference is that the passivation aid TBPS-454 (di-tert-butyl polysulfide) is not used.

[0053] Comparative Example 2

[0054] The sulfurized FHUDS-8 and FC-24 catalysts were loaded into the hydrogenation reactor at a volume ratio of 5:6. The first stage of catalyst passivation was carried out, and the passivated feedstock oil, passivation aid (TBPS-454, di-tert-butyl polysulfide) and hydrogen entered the hydrogenation reactor. The passivation aid accounted for 13% of the passivated feedstock oil mass. The reaction temperature was adjusted to 345°C in the hydrogenation pretreatment zone and 355°C in the hydrocracking zone. The volume space velocity was controlled to 0.55h -1 After the constant temperature operation, the catalyst was deactivated in the second stage. The amount of deactivator was adjusted to 24% of the deactivator mass of the deactivator raw oil. Other conditions remained unchanged and the constant temperature was continued for 26 hours.

[0055] After the passivation constant temperature is completed, all the passivated raw oil is directly switched to catalytic diesel. Based on the weight of catalytic diesel, the passivator content is reduced to 12%, and the addition of passivator is stopped to carry out subsequent hydrofining and hydrocracking reactions.

[0056] The passivated stock oil and stock oil used in the examples and comparative examples are shown in Table 1. The main operating process conditions and gasoline product properties in the examples and comparative examples of the present invention are shown in Table 2.

[0057] Table 1 Properties of passivated raw materials and catalytic diesel

[0058] Passivation crude oil Catalytic diesel <![CDATA[Density (20 °C), g / cm 3 > 0.8565 0.9509 Distillation range, ℃ (ASTMD86) Initial distillation point / 10% 205 / 269 189 / 132 50％ / 90％ 286 / 355 286 / 356 95% / dry point (final distillation point) 369 / 378 371 / 401 Sulfur, μg / g 9800 3235 Nitrogen, μg / g 98 421 Aromatic content, wt% 24.5 69 PAHs content, wt% - 48

[0059] Table 2 Main operating process conditions and gasoline product properties

[0060]

[0061]

[0062] It can be seen from the data results of Examples 1-4 and Comparative Examples 1-2 that the method of the present invention, by optimizing the passivation raw materials, formulating the passivation aid, and the gradient passivation steps, effectively controls the catalyst carbon deposition rate, realizes the passivation process of the enhanced catalyst, and greatly shortens the gasoline product qualification cycle.

Claims

1. A method for producing gasoline by catalytic hydrogenation of diesel, characterized in that: include: (1) The passivated feedstock oil, the passivation aid and the hydrogen enter the hydrogenation reactor and sequentially contact the hydrogenation pretreatment catalyst and the hydrocracking catalyst to undergo a passivation reaction; (2) After the passivation reaction is completed, the passivation feedstock oil is switched to catalytic diesel in batches until the final feedstock is a mixture of catalytic diesel and passivation aid; specifically: based on the total weight of the passivation feedstock oil and catalytic diesel, the proportion of catalytic diesel is increased by 10% to 35% each time; based on the total weight of the passivation feedstock oil and catalytic diesel, the proportion of catalytic diesel is increased by 1% to 5% each time, and a constant temperature treatment is performed; Based on the total weight of the catalytic diesel in the final feed, the content of the passivation aid in the final feed is 5% to 20%; (3) When the switching process of step (2) is completed, the injection of the passivation agent is stopped, and the catalytic diesel and hydrogen are contacted with the passivated hydrogenation pretreatment catalyst and the hydrocracking catalyst in sequence to perform hydrogenation pretreatment reaction and hydrocracking reaction, respectively, to obtain gasoline; The initial distillation point of the passivated feedstock oil is 150°C to 350°C, and the final distillation point is 370°C to 420°C; The passivation aid is selected from one or more of di-tert-butyl sulfide, tert-nonyl polysulfide, tert-dodecyl polysulfide, diphenyl sulfide, and di-tert-nonyl sulfide; The passivation reaction includes two stages. In the first stage, the passivation aid accounts for 1% to 15% of the passivation raw oil mass, and the operation time is 16 to 72 hours; in the second stage, the passivation aid accounts for 16% to 30% of the passivation raw oil mass, and the operation time is 12 to 48 hours.

2. The method according to claim 1, characterized in that The passivation aid is selected from one or more of di-tert-butyl polysulfide and diphenyl polysulfide.

3. The method according to claim 1 or 2, characterized in that: The passivation aid is selected from one or more of di-tert-butyl disulfide and diphenyl disulfide.

4. The method according to claim 1, characterized in that The initial distillation point of the passivated feedstock oil is 150° C. to 250° C., the aromatics content is less than 30wt%, and the nitrogen content is less than 200mg / g.

5. The method according to claim 4, characterized in that The passivated feedstock oil has an aromatic content of 15wt% to 25wt% and a nitrogen content of 50 to 100mg / g.

6. The method according to claim 1, characterized in that The conditions of the passivation reaction are: reaction temperature is 310℃~380℃, volume space velocity is 0.2h -1 ~6.0h -1 The volume ratio of hydrogen to oil is 200:1~2000:1, and the total passivation time is 40~100h.

7. The method according to claim 6, characterized in that The conditions of the passivation reaction are: reaction temperature is 320℃~350℃, volume space velocity is 0.6h -1 ~5h -1 , the volume ratio of hydrogen to oil is 500:1~1500:

1.

8. The method according to claim 1, characterized in that: The passivation reaction includes two stages. In the first stage, the operation time is 24 to 48 hours. In the second stage, the passivation additive accounts for 20% to 30% of the mass of the passivation raw oil, and the operation time is 16 to 32 hours.

9. The method according to claim 1, characterized in that: The initial distillation point of the catalytic diesel is 65°C-290°C, the final distillation point is 330°C-450°C, the aromatic content of the catalytic diesel fraction is ≥50wt%, and the polycyclic aromatic hydrocarbon content is ≥35wt%.

10. The method according to claim 9, characterized in that The initial boiling point of the catalytic diesel is 100-200° C., the aromatic content of the catalytic diesel fraction is 50wt%-75wt%, and the polycyclic aromatic hydrocarbon content is 40wt%-50wt%.

11. The method according to claim 1, characterized in that: The conditions of the hydrogenation pretreatment reaction include: reaction temperature of 260°C to 430°C, volume space velocity of 0.4h -1 ~10.0h -1 , hydrogen-oil volume ratio is 200:1~2000:1, total pressure is 2.5~20.0MPa; and / or, The conditions of the hydrocracking reaction include: a reaction temperature of 350°C to 450°C, a volume space velocity of 0.4h -1 ~16.0h -1 The volume ratio of hydrogen to oil is 200:1~2000:1, and the total pressure is 2.5~20.0MPa.

12. The method according to claim 11, characterized in that The conditions of the hydrogenation pretreatment reaction include: reaction temperature of 300°C to 380°C, volume space velocity of 0.6h -1 ~5h -1 , hydrogen-oil volume ratio is 500:1~1500:1, total pressure is 5.0~15.0MPa; and / or, The conditions of the hydrocracking reaction include: a reaction temperature of 355°C to 420°C, a volume space velocity of 0.6 h -1 ~6h -1 The volume ratio of hydrogen to oil is 500:1~1500:1, and the total pressure is 5.0~15.0MPa.

13. The method according to claim 1, characterized in that The sulfur content of the obtained gasoline product is ≯10mg / g, and the research octane number is ≮92.

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