A method for hydrotreating coking gasoline and diesel

Abstract

The present invention discloses a method for hydrotreating coked gasoline and diesel. The method comprises: cutting the coked gasoline and diesel feedstock into a heavy fraction and a light fraction; dividing the reaction zone into a hydrogenation reaction zone I loaded with a hydrorefining catalyst and a hydrogenation reaction zone II loaded with a desiliconization and deolefination catalyst; mixing the heavy fraction with hydrogen and passing it from bottom to top through the hydrogenation reaction zone I; mixing the reaction stream from the hydrogenation reaction zone I with the light fraction, a desiliconization aid, and hydrogen and passing it from bottom to top through the hydrogenation reaction zone II; and the reaction stream from the hydrogenation reaction zone II comprising a liquid stream and a gaseous stream. A portion of the liquid stream is recycled and enters the hydrogenation reaction zone I together with the heavy fraction and hydrogen; the remaining portion of the liquid stream and the entire gaseous stream enter a subsequent separation system to separate and obtain gasoline and diesel products. The method reduces hydrogen consumption, increases liquid yield, and can produce high-quality diesel and gasoline products for ethylene cracking.

Description

Technical Field

[0001] The invention belongs to the field of petroleum refining and relates to a coking gasoline and diesel hydrogenation treatment method. Background Art

[0002] Delayed coking is a key method for refining and petrochemical companies to process low-quality heavy oil, but delayed coking products generally require further impurity removal. For example, coking gasoline and diesel contain high levels of impurities such as sulfur, nitrogen, olefins, gums, and silicon-containing compounds. The gasoline fraction is used as a feedstock for ethylene cracking, fertilizers, and reforming units. Jet kerosene and diesel fractions can produce high-quality fuel oil, but both require hydrorefining and desiliconization. Silicon-containing compounds are derived from the silicon-containing defoamers added during the delayed coking process. The silicon in these defoamers can enter the reactor with the feedstock and deposit on the catalyst, causing silicon poisoning and permanent deactivation of the catalyst, which can also shorten the unit's single-cycle run. Several publications, such as "Direct Determination of Silicon Content in Coking Distillates by ICP-OES," report that the primary product of thermal cracking of polydimethylsiloxane as a defoamer is cyclosiloxane, primarily found in the naphtha fraction. The optimal reaction temperature for desiliconization is above 270°C for silicon-capturing catalysts, which are typically used in combination with hydrodesulfurization catalysts.

[0003] The article "Research on Hydrogenation of Coking Gasoline and Diesel to Produce Jet Fuel" published in Petrochemical Technology and Application, Issue 3, 2004, pp. 173-176, describes deep hydrorefining of a mixture of coking gasoline and diesel, followed by fractionation into gasoline, kerosene, and diesel. This method only deeply refines the full coking fraction and then separates the kerosene fraction, resulting in excessive hydrogenation. Sun Guang's article "Low-Pressure Hydrorefining of Coking Gasoline and Diesel Using FH-5 Catalyst" published in Refining Technology and Engineering, Issue 5, 1994, pp. 39-42, describes the use of a hydrorefining unit to hydrorefining coking gasoline and coking diesel. The method employs a feed switching method, whereby a batch of coking gasoline is processed and then the coking diesel feedstock is switched. This method is only suitable for small-scale coking units, requiring constant switching and frequent changes in operating conditions during the production process, which is inconvenient and significantly affects catalyst performance.

[0004] CN101003751A discloses a method for processing coker full-fraction oil, comprising separating the coker full-fraction oil into coker light fraction oil and coker heavy fraction oil, wherein the light fraction oil contains a portion of light diesel oil fraction, and the heavy fraction oil is a heavy diesel oil fraction with a higher dry weight. The coker heavy fraction oil is hydrocracking, and the cracking products are mixed with the coker light fraction oil and then subjected to hydrorefining treatment. This method suffers from problems such as over-hydrogenation of the light fraction and over-desulfurization.

[0005] CN112852480A discloses a grading method for a coking gasoline hydrorefining catalyst and a hydrorefining method, comprising: sequentially arranging a replaceable zone and a main reaction zone in series along the direction of logistics, wherein the replaceable zone is sequentially loaded with a diolefin saturation catalyst in a first reactor and a silicon scavenger in a second reactor along the direction of logistics, wherein the second reactor is a parallel dual reactor capable of online switching; and the main reaction zone is sequentially loaded with an arsenic removal agent and a gasoline hydrorefining catalyst along the direction of logistics.

[0006] CN111321005A discloses a low-energy, long-cycle hydrogenation process for producing diesel. The method first fractionates the diesel feedstock into a light fraction and a heavy fraction, passes the heavy fraction through a first reaction zone for a hydrodesulfurization reaction, and then reacts with the light fraction diesel feedstock in a second hydrogenation reaction zone under the action of a hydrodesulfurization catalyst.

[0007] The processes used in the above-mentioned existing technologies all rely on conventional gas-phase circulating trickle-bed hydrogenation. However, these processes fail to consider the exothermic control of the deolefination reaction and the matching of the desiliconization catalyst with the desulfurization requirements of the reaction fraction. Some products suffer from excessive hydrodesulfurization, which can lead to problems such as excessive product quality, reduced liquid yield, and excessive hydrogen and energy consumption. Summary of the Invention

[0008] In view of the shortcomings of the prior art, the present invention aims to provide a method for hydrotreating coking gasoline and diesel, which can better deeply couple the reaction fractions and the reaction temperature, that is, realize reasonable desulfurization, desiliconization and deolefination reactions, and avoid excessive hydrodesulfurization reactions of some fractions, thereby reducing hydrogen consumption and improving liquid yield. At the same time, the reaction severity is reduced, the role of each reaction zone can be better exerted, the heat balance can be better utilized, the energy consumption of the device can be reduced, and high-quality diesel and gasoline products for ethylene cracking can be produced.

[0009] The coking gasoline and diesel hydroprocessing method of the present invention comprises: cutting the coking gasoline and diesel raw materials into heavy fractions and light fractions, dividing the reaction zone into a hydrogenation reaction zone I filled with a hydrorefining catalyst and a hydrogenation reaction zone II filled with a desiliconization and deolefination catalyst, mixing the heavy fraction with hydrogen and passing it from bottom to top through the hydrogenation reaction zone I, mixing the reaction stream obtained in the hydrogenation reaction zone I with the light fraction, a desiliconization auxiliary agent, and hydrogen and passing it from bottom to top through the hydrogenation reaction zone II, the reaction stream obtained in the hydrogenation reaction zone II comprises a liquid phase stream and a gas phase stream, a portion of the liquid phase stream is recycled and enters the hydrogenation reaction zone I together with the heavy fraction and hydrogen, and the remaining portion of the liquid phase stream and the entire gas phase stream enter a subsequent separation system to separate and obtain gasoline and diesel products.

[0010] Furthermore, the initial boiling point temperature of the coking gasoline and diesel raw materials is 30°C to 110°C, and the final boiling point temperature is 350°C to 410°C. The silicon content in the coking gasoline and diesel raw materials is 2 to 500 μg.g-1 , preferably 5 to 100 μg.g -1 , olefin content 5wt%~40wt%, sulfur content>5000μg.g -1 , preferably 6000~15000μg.g -1 , nitrogen content ≤ 2000μg.g -1 The coking gasoline and diesel raw materials contain more than 50 wt% of coking gasoline and diesel, and can be mixed and processed into one or more of straight-run diesel, diesel derived from hydrogenation of residual oil, catalytic gasoline, ethylene cracking gasoline, straight-run naphtha and other oil products.

[0011] Furthermore, the cutting point for cutting the coking gasoline raw material into heavy fraction and light fraction is 160°C to 250°C, wherein the proportion of the heavy fraction is not less than 35wt%, preferably 40wt% to 70wt% based on the mass of the coking gasoline and diesel raw material.

[0012] Furthermore, the hydrorefining catalyst generally uses refractory porous oxides as carriers, such as aluminum oxide, silicon oxide, amorphous silicon aluminum, titanium oxide, and composite oxides or mixed oxide carriers of several elements. Generally, non-acidic or weakly acidic materials are used as carriers, and the hydrogenation activity combination of the hydrorefining catalyst is W-Ni, Mo-Ni or W-Mo-Ni. The content of hydrogenation active components in terms of oxides is generally 15wt% to 60wt%, preferably 16wt% to 46wt%. Among them, the nickel oxide content is 1.8wt% to 8.5wt%, preferably 2.0wt% to 5.5wt%. The shape is spherical or bar-shaped, the spherical diameter is 0.1 to 12mm, preferably 0.4 to 5mm; the bar is 2 to 15mm in length, preferably 2 to 8mm, and the diameter is 1 to 6mm, preferably 1.2 to 3.5mm. The specific surface area is 150 to 650m 2 / g, preferably 200-450m 2 / g, and a pore volume of 0.2-1.5 mL / g, preferably 0.30-0.85 mL / g. The hydrorefining catalyst can be a commercial catalyst selected according to the process requirements, or can be prepared according to existing methods. It can also be a regenerated catalyst obtained by regenerating a deactivated catalyst, such as FHUDS-10 produced by the Fushun Branch of Sinopec Catalyst Company.

[0013] Furthermore, the active metals of the desiliconization and deolefination catalyst are Group VIB metals and Group VIII metals. Based on the total weight of the catalyst, the Group VIB metals are calculated as oxides at 4% to 45%, preferably 10% to 35%, and the Group VIII metals are calculated as oxides at 2% to 20%, preferably 4% to 15%. The pore volume of the desiliconization and deolefination catalyst is 0.5 to 1.0 mL / g, preferably 0.7 to 1.0 mL / g; the specific surface area is 260 to 550 m2 / g, preferably 350 to 500 m 2 / g; its shape is spherical or bar-shaped, with a spherical diameter of 0.2-20 mm, preferably 0.3-5 mm; the bar length is 2-15 mm, preferably 3-8 mm, and the diameter is 0.6-6 mm, preferably 0.8-3.5 mm. The catalyst preparation method is well known in the art and can adopt one or more of the following methods: impregnation, co-extrusion, and co-precipitation, such as FHRS-3 produced by Sinopec Catalyst Company Fushun Branch.

[0014] Furthermore, the volume ratio of the hydrorefining catalyst loaded in the hydrogenation reaction zone I to the hydrodesiliconization and deolefination catalyst loaded in the hydrogenation reaction zone II is 80:20 to 40:60, preferably 70:30 to 50:50.

[0015] Furthermore, the hydrogenation reaction zone I and the hydrogenation reaction zone II can be placed in different reactors or in the same reactor, wherein the hydrogenation reaction zone I adopts a fixed bed and the hydrogenation reaction zone II adopts an ebullating bed.

[0016] Furthermore, the reaction conditions of the hydrogenation reaction zone I include: hydrogen partial pressure of 4.0 MPa to 16.0 MPa, preferably 8.0 MPa to 12.0 MPa, volume space velocity of 0.3 h -1 ~10.0h -1 The volume ratio of hydrogen to oil is 150:1 to 1000:1, and the average reaction temperature is 340°C to 420°C, preferably 350°C to 410°C.

[0017] Furthermore, the reaction conditions of the hydrogenation reaction zone II include: hydrogen partial pressure of 1.0 MPa to 12.0 MPa, preferably 3.0 MPa to 8.0 MPa, volume space velocity of 0.8 h -1 ~10.0h -1 The volume ratio of hydrogen to oil is 200:1 to 2500:1, and the average reaction temperature is 250°C to 320°C, preferably 270°C to 300°C.

[0018] Furthermore, the average reaction temperature of the hydrogenation reaction zone I is 50° C. to 140° C. higher than the average reaction temperature of the hydrogenation reaction zone II. The hydrogen partial pressure of the hydrogenation reaction zone I is 2.0 MPa to 8.0 MPa higher than the hydrogen partial pressure of the hydrogenation reaction zone II.

[0019] Furthermore, the desiliconization aid is selected from alcohols and / or organic acids, preferably selected from C1 to C10 alcohols and / or organic acids, and further preferably selected from C2 to C6 alcohols and / or organic acids, for example, one or more of ethanol, ethylene glycol, propanol, propylene glycol, glycerol, butanol, butanediol, butanetriol, pentanol, acetic acid, oxalic acid, tartaric acid, succinic acid, hydroxyacetic acid and citric acid.

[0020] Furthermore, based on the mass of the light fraction, the mass content of the desiliconization aid is 0.01 wt% to 5 wt%, preferably 0.5 wt% to 4.5 wt%, and more preferably 2.0 wt% to 4.0 wt%.

[0021] Furthermore, the volume ratio of the recycled portion of the liquid phase logistics to the total liquid phase logistics is 0.3 to 0.7:1.

[0022] Furthermore, the separation system includes a cold high fractionator, a cold low fractionator and a fractionation unit.

[0023] Furthermore, the remaining liquid phase and the entire gas phase flow first enter the cold high-pressure fractionation unit. The separated circulating hydrogen is desulfurized (using conventional desulfurization methods in the hydrogenation industry) and then mixed with fresh hydrogen through a circulating hydrogen compressor. The liquid phase separated by the cold high-pressure fractionation unit enters the cold low-pressure fractionation unit and then enters the fractionation unit to produce gasoline and diesel products. The fractionation unit is capable of precise gasoline and diesel fractionation.

[0024] The hydrotreating method of coking gasoline and diesel of the present invention has the following beneficial effects:

[0025] 1. The coking gasoline and diesel hydroprocessing method provided by the present invention divides the coking gasoline and diesel raw materials into different fractions, optimizes the catalyst system and reaction form according to different product requirements, and deeply couples the reaction fraction, reaction sequence, reaction temperature, reaction pressure and various reactions to achieve precise control of reaction depth, balance of reaction heat, and full utilization of catalyst performance, thereby achieving deep desulfurization of the diesel fraction, controlling the strong exothermicity of olefins in the gasoline fraction, and achieving efficient desiliconization and deolefination of the light fraction with the help of a desiliconization additive, while avoiding excessive desulfurization, thereby achieving the purpose of cascade control of the reaction process.

[0026] 2. The present invention vaporizes the light fraction and conducts an efficient desiliconization and deolefination reaction with the liquid phase heavy fraction after the hydrodesulfurization reaction in the upper reaction zone. Since it is an ebullated bed reaction, the light fraction vaporizes and enters the ebullated bed reaction zone in an approximately isothermal reaction. As the gas phase contacts the boiling catalyst upward, the reaction is carried out, thereby controlling the isothermal reaction temperature of the desiliconization and deolefination reaction zone, achieving the matching of the light fraction, desiliconization adjuvant, desiliconization and deolefination catalyst activity and reactants, and effectively changing the mismatch of the desiliconization catalytic reaction temperature and the traditional passive desulfurization. According to the inherent characteristics of the ebullated bed reaction, the reaction temperature of the desiliconization and deolefination catalyst bed is effectively controlled, which effectively solves the problem of reaction exotherm in the original process and avoids the problem of reaction temperature mismatch in the initial operation of the silicon-capturing catalyst. At the same time, the ebullated bed reaction form can add and discharge the catalyst online, avoiding silicon deposition in the catalyst, which causes the catalyst system activity to decrease, and is more conducive to extending the operation cycle.

[0027] 3. The two reaction zones provided by the present invention are respectively loaded with a hydrorefining catalyst and a desiliconization and deolefination catalyst, and the corresponding reaction temperatures and reaction pressures are adjusted accordingly, following the order of reactants from difficult to easy, and the reaction temperatures and reaction pressures from high to low. This effectively maximizes the catalyst's high-temperature hydrodesulfurization, temperature-controlled desiliconization, and deolefination performance, avoiding the passive over-desulfurization of sulfur in the gasoline fraction to ensure ultra-deep removal of thiophenic sulfur in conventional coking gasoline and diesel hydrogenation processes. In particular, the desiliconization reaction temperature is effectively controlled in the ebullient bed reaction to avoid excessive desulfurization of the coker gasoline fraction. This reduces the amount of sulfur injected into the downstream ethylene unit, improves total liquid yield, increases the yield of gasoline components, and reduces hydrogen consumption.

[0028] 4. The method of the present invention adopts gradient reaction heat release and effectively matches heat capacity and reaction heat balance, effectively controls reaction temperature rise, reduces the probability of thermal cracking, ensures high liquid yield, and low hydrogen consumption and energy consumption.

[0029] 5. The method of the present invention can be used to slightly modify existing coking gasoline and diesel hydrogenation or other similar hydrogenation equipment, which reduces the modification cost. At the same time, the operation steps are reduced. The process is simple, easy to operate, safe, environmentally friendly, and has a long operating cycle. BRIEF DESCRIPTION OF THE DRAWINGS

[0030] Figure 1 is a schematic diagram of the process adopted in the embodiment of the present invention;

[0031] Among them, 1-heavy fraction, 2-light fraction, 5-hydrogenation reaction zone I, 6-hydrogenation reaction zone II, 7-gas phase logistics, 8-liquid phase logistics, 10-cold high fraction, 12-cold low fraction, 15-circulating hydrogen compressor. DETAILED DESCRIPTION

[0032] The following non-limiting examples 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.

[0033] The following combination Figure 1 The method disclosed in the present invention is described in more detail with reference to the accompanying drawings and specific examples.

[0034] like Figure 1As shown, coking gasoline and diesel are cut into light and heavy fractions. The heavy fraction 1 is mixed with the circulating hydrogen in pipeline 16 and supplemented with new hydrogen according to actual conditions. It passes through pipeline 4 and enters the fixed-bed hydrogenation reaction zone I from top to bottom for hydrogenation reactions such as hydrodesulfurization and hydrodenitrogenation. The reaction products are mixed with the light fraction 2 and the circulating hydrogen in pipeline 17 and enter the ebullating bed hydrogenation reaction zone II from top to bottom for reactions such as hydrodesiliconization, hydrodeolefination and hydrodesulfurization. Liquid stream 8 flows out from the top of the reactor, part of which enters the bottom of the reaction zone I and is recycled back to the reaction zone I through pipeline 9; gaseous stream 7 flows out from the top of the reactor, mixes with the remaining liquid reactants, and enters the cold high fraction 10. The separated liquid phase enters the cold low fraction 12 through pipeline 11 and then enters the fractionation unit through pipeline 13; the circulating hydrogen separated in the cold high fraction 10 enters the circulating hydrogen desulfurization system through pipeline 14, then passes through the circulating hydrogen compressor 15, returns to pipeline 16, and is mixed with the new hydrogen passing through pipeline 2. Figure 1 Many necessary equipment are omitted, such as oil pumps, valves, heating furnaces, etc.

[0035] The hydrorefining catalyst used in the Examples and Comparative Examples is FHUDS-10, developed by the Fushun Petrochemical Research Institute and produced by the Fushun Branch of Sinopec Catalyst Company. The desiliconization and deolefination catalyst used in the Examples and Comparative Examples is FHRS-3, developed by the Fushun Petrochemical Research Institute and produced by the Fushun Branch of Sinopec Catalyst Company. The FHUDS-10 catalyst uses an alumina support and Mo-Ni as the active metal component. It is in the form of a bar, 2 to 8 mm long and 1.2 to 1.5 mm in diameter. The FHRS-3 catalyst also uses an alumina support and Mo-Ni as the active metal component. It is in the form of a bar, 3 to 5 mm long and 1.0 to 3.5 mm in diameter.

[0036] Example 1

[0037] use Figure 1 The hydrogenation process flowsheet describes a process where the coked gasoline and diesel feedstock is cut at 230°C into a light fraction (45 wt%) and a heavy fraction (55 wt%). The heavy fraction is then mixed with hydrogen and passed from bottom to top through hydrogenation zone I (loaded with a hydrorefining catalyst). The resulting reaction stream from hydrogenation zone I is mixed with the light fraction, ethylene glycol (1.5% by weight based on the mass of the light fraction), and hydrogen, and passed from bottom to top through hydrogenation zone II (loaded with a desilication and deolefination catalyst). The resulting reaction stream from hydrogenation zone II consists of a liquid phase and a vapor phase. Part of the liquid phase is recycled at a volume ratio of 0.4:1. The recycled liquid phase enters hydrogenation zone I along with the heavy fraction and hydrogen. The remaining liquid phase and the entire vapor phase enter a subsequent separation system to separate the gasoline and diesel products. The properties of the coked gasoline and diesel are shown in the table. Specific process conditions and product properties are shown in Table 2.

[0038] Example 2

[0039] Compared with Example 1, the difference is that the desiliconization auxiliary agent used is citric acid. The properties of coking gasoline and diesel are shown in the table. The specific process conditions and product properties are shown in Table 2.

[0040] Example 3

[0041] Compared with Example 1, the difference is that ethylene glycol, a desiliconizing agent, is added. Based on the mass of the light fraction, the mass content of the desiliconizing agent is 2.0%. The properties of the coking gasoline and diesel are shown in the table. The specific process conditions and product properties are shown in Table 2.

[0042] Example 4

[0043] Compared with Example 1, the difference is that citric acid is added as a desiliconizing agent. Based on the mass of the light fraction, the mass content of the desiliconizing agent is 2.0%. The properties of the coking gasoline and diesel are shown in the table. The specific process conditions and product properties are shown in Table 2.

[0044] Comparative Example 1

[0045] The conventional coking gasoline and diesel hydrorefining process is adopted. The reactor is a trickle bed reactor. The coking gasoline and diesel are mixed with hydrogen and enter the trickle bed reactor from top to bottom. They first come into contact with the desiliconization catalyst bed filled with FHRS-3, and then with the hydrodesulfurization catalyst bed filled with FHUDS-10, to carry out hydrodeolefination, desiliconization and desulfurization reactions. The volume ratio of the hydrodesulfurization catalyst to the desiliconization catalyst is 60:40. The volume ratio of hydrogen to oil is 450:1, and the volume space velocity is 1.0h -1 The average reaction temperature of the desiliconization catalyst bed was 328°C, and the average reaction temperature of the hydrodesulfurization catalyst bed was 365°C. Specific process conditions and product properties are shown in Table 2.

[0046] Table 1 Properties of coking gasoline and diesel feedstock

[0047] Coking gasoline and diesel <![CDATA[Density (20 °C), g / cm 3 > 0.8255 Distillation range, ℃ (ASTMD86) 50～376 Sulfur, μg / g 9615 Nitrogen, μg / g 565 Olefins, wt% 15 <![CDATA[Silicon content, μg.g -1 > 6

[0048] Table 2 Process conditions and properties of gasoline and diesel products

[0049]

[0050]

[0051] Table 2

[0052]

[0053] As can be seen from the data results of the Examples and Comparative Examples, the method of the present invention utilizes a simple modification of a conventional coking gasoline and diesel hydrorefining unit, combining the advantages of fixed-bed and ebullating-bed reaction modes. Furthermore, by optimizing catalyst gradation, fraction cutting, and incorporating the reaction requirements of the reactants, a deep coupling of reaction phase, reaction mode, and reaction process conditions is achieved. This controls the reaction depth, balances the reaction heat, maximizes the performance of the catalysts in each reaction zone, achieves efficient desulfurization while avoiding over-desulfurization, and reduces hydrogen and energy consumption. Furthermore, product liquid yield is increased, effectively extending the operating cycle of the coking gasoline hydrorefining unit.

Claims

1. A method for hydrotreating coking gasoline and diesel, comprising: The coking gasoline and diesel feedstock is cut into a heavy fraction and a light fraction, and the reaction area is divided into a hydrogenation reaction zone I filled with a hydrorefining catalyst and a hydrogenation reaction zone II filled with a desiliconization and deolefination catalyst. The heavy fraction is mixed with hydrogen and passes through the hydrogenation reaction zone I from bottom to top. The reaction stream obtained in the hydrogenation reaction zone I is mixed with the light fraction, a desiliconization auxiliary agent, and hydrogen and passes through the hydrogenation reaction zone II from bottom to top. The reaction stream obtained in the hydrogenation reaction zone II includes a liquid stream and a gas stream. A portion of the liquid stream is recycled and enters the hydrogenation reaction zone I together with the heavy fraction and hydrogen. The remaining portion of the liquid stream and the entire gas stream enter a subsequent separation system to separate and obtain gasoline and diesel products. The hydrogenation reaction zone I is a fixed bed hydrogenation reaction, and the hydrogenation reaction zone II is an ebullient bed hydrogenation reaction; The average reaction temperature of the hydrogenation reaction zone I is 50°C to 140°C higher than the average reaction temperature of the hydrogenation reaction zone II; the hydrogen partial pressure of the hydrogenation reaction zone I is 2.0MPa to 8.0MPa higher than the hydrogen partial pressure of the hydrogenation reaction zone II; The desiliconization aid is selected from alcohols and / or organic acids.

2. The method according to claim 1, characterized in that The initial boiling point temperature of the coking gasoline and diesel raw materials is 30℃~110℃, and the final boiling point temperature is 350℃~410℃; the silicon content in the coking gasoline and diesel raw materials is 2~500μg.g -1 , olefin content 5wt%~40wt%, sulfur content ≤15000μg.g -1 , nitrogen content ≤ 2000μg.g -1 .

3. The method according to claim 2, characterized in that The silicon content in the coking gasoline and diesel raw materials is 5 to 100 μg.g -1 .

4. The method according to claim 1, wherein The cutting point for cutting the coking gasoline raw material into heavy fraction and light fraction is 160° C. to 250° C., wherein the proportion of the heavy fraction is not less than 35wt% based on the mass of the coking gasoline and diesel raw material.

5. The method according to claim 4, characterized in that Based on the mass of the coking gasoline and diesel raw materials, the proportion of the heavy fraction is 40wt% to 70wt%.

6. The method according to claim 1, wherein The volume ratio of the hydrorefining catalyst filled in the hydrogenation reaction zone I to the hydrodesiliconization and deolefination catalyst filled in the hydrogenation reaction zone II is 80:20 to 40:

60.

7. The method according to claim 6, characterized in that The volume ratio of the hydrorefining catalyst filled in the hydrogenation reaction zone I to the hydrodesiliconization and deolefination catalyst filled in the hydrogenation reaction zone II is 70:30 to 50:

50.

8. The method according to claim 1, characterized in that The reaction conditions of the hydrogenation reaction zone I include: hydrogen partial pressure of 4.0 MPa to 16.0 MPa, volume space velocity of 0.3 h -1 ~10.0h -1 The volume ratio of hydrogen to oil is 150:1 to 1000:1, and the average reaction temperature is 340℃ to 420℃.

9. The method according to claim 8, characterized in that The reaction conditions of the hydrogenation reaction zone I include: hydrogen partial pressure of 8.0 MPa to 12.0 MPa, and average reaction temperature of 350° C. to 410° C.

10. The method according to claim 1, characterized in that The reaction conditions of the hydrogenation reaction zone II include: hydrogen partial pressure of 1.0 MPa to 12.0 MPa, volume space velocity of 0.8 h -1 ~10.0h -1 The volume ratio of hydrogen to oil is 200:1 to 2500:1, and the average reaction temperature is 250℃ to 320℃.

11. The method according to claim 10, characterized in that The reaction conditions of the hydrogenation reaction zone II include: a hydrogen partial pressure of 3.0 MPa to 8.0 MPa, and an average reaction temperature of 270° C. to 300° C.

12. The method according to claim 1, characterized in that The desiliconization aid is selected from C1-C10 alcohols and / or organic acids.

13. The method according to claim 12, characterized in that The desiliconization aid is selected from C2-C6 alcohols and / or organic acids.

14. The method according to claim 12, characterized in that The desiliconization auxiliary agent is one or more of ethanol, ethylene glycol, propanol, propylene glycol, glycerol, butanol, butanediol, butanetriol, pentanol, acetic acid, oxalic acid, tartaric acid, succinic acid, glycolic acid and citric acid.

15. The method according to claim 1, wherein Based on the mass of the light fraction, the mass content of the desiliconization auxiliary agent is 0.01wt% to 5wt%.

16. The method according to claim 15, characterized in that Based on the mass of the light fraction, the mass content of the desiliconization auxiliary agent is 0.5wt% to 4.5wt%.

17. The method according to claim 15, characterized in that Based on the mass of the light fraction, the mass content of the desiliconization auxiliary agent is 2.0wt% to 4.0wt%.

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