A coking gasoline hydrogenation process

Abstract

The present invention discloses a coker gasoline hydrogenation process, comprising the following steps: a coker gasoline feedstock is cut into a light fraction and a heavy fraction; the heavy fraction is mixed with hydrogen and enters a hydrodesulfurization reaction zone, where it contacts a hydrodesulfurization catalyst for a hydrodesulfurization reaction; the hydrodesulfurization reaction effluent is mixed with the light fraction and enters a desiliconization and deolefination reaction zone, where it contacts a silicon scavenger and a hydrodeolefination catalyst for desiliconization and deolefination reactions; the reaction product undergoes gas-liquid separation and enters a fractionation zone, where it is stripped and fractionated to obtain an ethylene cracking material. The process of the present invention deeply and organically combines the reaction fractions, reaction sequence, and various reaction types, controls the reaction depth, balances the reaction heat, maximizes the efficiency of the catalyst in the reaction, reduces device energy consumption, and extends the operating cycle of the coker gasoline hydrogenation device.

Description

Technical Field

[0001] The invention belongs to the field of petroleum refining and relates to a hydrogenation process for coking gasoline, in particular to a hydrogenation process for producing ethylene cracking material by utilizing coking gasoline. Background Art

[0002] Delayed coking is one of the important means for refining companies to process low-quality heavy oil, but delayed coking products generally require further impurity removal. For example, coking naphtha contains high levels of impurities such as sulfur, nitrogen, olefins, gums, and silicon-containing compounds. If it is to be used as a raw material for units such as ethylene cracking, fertilizers, and reforming, it must first be hydrorefined and desiliconized. The silicon-containing compounds come from the silicon-containing defoaming agents added during the delayed coking process. The silicon in the silicon-containing defoaming agent will enter the reactor with the raw oil and deposit on the catalyst, causing the catalyst to be permanently poisoned and deactivated by silicon, and also resulting in a short single start-up cycle of the unit. Many documents such as "Direct Determination of Silicon Content in Coking Distillate Oil by ICP-OES" report that the main product generated after the thermal cracking of polydimethylsiloxane as a defoaming agent is cyclosiloxane.

[0003] Currently, the coker gasoline hydrogenation process involves a simple, one-step desiliconization, hydrodeolefinization, and desulfurization process. The coking reaction is a thermal cracking process, and the coker oil contains a high concentration of olefins. Conventional coker gasoline hydrogenation processes require that the reactants, upon entering the reactor, first react with a protective agent and a silicon scavenger to remove impurities and silicon, followed by hydrodeolefination and desulfurization, allowing the silicon scavenger to protect the main catalyst. However, since olefin saturation is more likely to occur than desiliconization and desulfurization, and the reactions are highly exothermic, coker gasoline hydrogenation can only be performed at a relatively low inlet reaction temperature to avoid the strong exothermic reaction of olefins and maintain a controllable reaction process. While the optimal reaction temperature for silicon scavenging catalysts is above 270°C, the initial temperature of coker gasoline hydrogenation is only around 200°C. This results in the silicon scavenger being unable to meet the initial desiliconization requirements, leading to premature silicon poisoning of the main catalyst and a rapid decrease in activity, impacting the operating cycle.

[0004] CN103789020A discloses a coker gasoline hydrogenation method. Combining the advantages of both an ebullating bed and a fixed bed, circulating oil is used to dilute unsaturated hydrocarbons such as diolefins in the feedstock while simultaneously removing the heat of reaction. This prevents coking of the catalyst bed and reduces the reactor pressure differential. This method requires no major modifications to existing equipment and can achieve long-term operation with minimal investment, thereby increasing economic benefits.

[0005] CN112852480A discloses a coker gasoline hydrogenation catalyst grading and hydroprocessing method, comprising: a replaceable zone and a main reaction zone arranged in series along the logistics direction. The replaceable zone is sequentially loaded with a diolefin saturation catalyst in the first reactor and a silicon scavenger in the second reactor, along the logistics direction. The second reactor is a parallel dual reactor capable of online switching. The main reaction zone is sequentially loaded with an arsenic removal agent and a gasoline hydrogenation catalyst along the logistics direction. This method can improve the impurity removal rate and extend the operating cycle of the gasoline hydrogenation unit. It is particularly suitable for gasoline hydrogenation units using coker gasoline as feedstock or with a high coker gasoline content.

[0006] CN107955642A discloses a method and system for hydrogenating a mixture of C5 raffinate oil and coker gasoline. The method comprises: subjecting the C5 raffinate oil to a pre-hydrogenation reaction to saturate the diolefins in the C5 raffinate oil with hydrogen; and then using the pre-hydrogenation product as a diluent and hydrogenating it together with the coker gasoline.

[0007] The processes used in the aforementioned patents are all conventional gas-phase circulating trickle-bed hydrogenation processes. They fail to integrate the exothermic heat of the coker gasoline reaction with the active temperature of the silicon scavenger. Because coker gasoline hydrogenation is highly exothermic, the initial reaction temperature of the silicon scavenger is often too low to match the activity of the silicon scavenger. Summary of the Invention

[0008] In response to the shortcomings of the prior art, the present invention provides a coker gasoline hydrogenation process. The process of the present invention deeply and organically combines the reaction fractions, reaction sequence and various reactions, controls the reaction depth, balances the reaction heat, maximizes the efficiency of the catalyst in the reaction, reduces the energy consumption of the device, and extends the operating cycle of the coker gasoline hydrogenation device.

[0009] The coker gasoline hydrogenation process of the present invention comprises the following steps: a coker gasoline raw material is cut into a light fraction and a heavy fraction, the heavy fraction is mixed with hydrogen and enters a hydrodesulfurization reaction zone, and contacts with a hydrodesulfurization catalyst to carry out a hydrodesulfurization reaction, the hydrodesulfurization reaction effluent is mixed with the light fraction and enters a desiliconization and deolefination reaction zone, and contacts with a silicon scavenger and a hydrodeolefination catalyst to carry out desiliconization and deolefination reactions, and the reaction product is subjected to gas-liquid separation and enters a fractionation zone, where it is stripped and fractionated to obtain an ethylene cracking material.

[0010] In the process of the present invention, the coker gasoline feedstock has an initial boiling point of 30-100°C, an end boiling point of 160-260°C, an olefin content of 15-50% by weight, a silicon content of 3-500µg / g, and a sulfur content of 5,000-15,000µg / g, preferably 6,000-12,000µg / g. At least 40-100% by weight of the coker gasoline feedstock is coker gasoline, and one or more of catalytic gasoline, ethylene cracking gasoline, and straight-run naphtha may also be mixed.

[0011] In the process of the present invention, the cutting point temperature is 135-170°C, the sulfur content in the light fraction after cutting is 3000µg / g-8000µg / g, and the sulfur content in the heavy fraction is 7000µg / g-18000µg / g.

[0012] In the process of the present invention, the hydrodesulfurization reaction zone is a trickle bed reaction process, and the hydrodesulfurization reaction process conditions are as follows: the inlet reaction temperature is 250-400°C, preferably 260-350°C, the hydrogen partial pressure is 1.0 MPa-16.0 MPa, preferably 3.0 MPa-10.0 MPa, and the volume space velocity is 0.3 h -1 ~16.0h -1 The volume ratio of hydrogen to oil is 100:1 to 2000:1.

[0013] In the process of the present invention, the concentration of hydrogen sulfide in the circulating hydrogen at the inlet of the hydrodesulfurization zone is 0 ppm to 2000 ppm, preferably 0 ppm to 1000 ppm.

[0014] In the process of the present invention, the concentration of hydrogen sulfide in the reaction effluent in the hydrodesulfurization zone is 5000 ppm to 20000 ppm, preferably 5000 ppm to 15000 ppm.

[0015] In the process of the present invention, the light fraction can be mixed with the product of the desulfurization reaction zone and enter the desiliconization and deolefination reaction zone; or the light fraction can be mixed with hydrogen and then mixed with the product of the desulfurization reaction zone and enter the desiliconization and deolefination reaction zone.

[0016] In the process of the present invention, the hydrodesulfurization catalyst is generally supported by refractory porous oxides, such as aluminum oxide, silicon oxide, titanium oxide, and composite oxides or mixed oxides of several elements. Generally, non-acidic or weakly acidic materials are used as carriers, and the hydrogenation activity of the catalyst is one or more of W, Mo, Ni and Co. The content of hydrogenation active components in terms of oxides is generally 3wt% to 55wt%, preferably 10wt% to 45wt%. The shape is spherical or bar-shaped, with a spherical diameter of 0.04 to 10mm, preferably 0.04 to 5mm; the bar is 2 to 10mm in length, preferably 2 to 8mm, and a diameter of 1 to 6mm, preferably 1.5 to 3.5mm. The pore volume of the hydrodesulfurization catalyst is 0.4-1.0mL / g, preferably 0.5-0.8mL / g; the specific surface area is 200-400m 2 / g, preferably 220-320m 2 The hydrodesulfurization catalyst can be a commercial catalyst selected according to the needs of the process, or can be prepared according to existing methods, or can be a regenerated catalyst obtained by regenerating a deactivated catalyst.

[0017] In the process of the present invention, the organic sulfur content in the hydrodesulfurization reaction effluent is 200µg / g to 800µg / g.

[0018] In the process of the present invention, the desiliconization and deolefination reaction zone is a trickle bed reaction process, and the reaction conditions are: the inlet reaction temperature is 270℃~350℃, preferably 285℃~300℃, the hydrogen partial pressure is 0.5MPa~16MPa, preferably 1MPa~6MPa, and the volume space velocity is 2h -1 ~6h -1 The volume ratio of hydrogen to oil is 100:1 to 2000:1.

[0019] In the process of the present invention, the volume ratio of the silicon-trapping agent and the hydrogenation deolefination catalyst in the desiliconization and deolefination reaction zone is 10:90 to 95:5, preferably 35:65 to 80:20; wherein the silicon-trapping agent is well known to those skilled in the art. As one of the specific embodiments, the active metals of the silicon-trapping catalyst are oxides of Group VIB metals and oxides of Group VIII metals. Based on the total weight of the catalyst, the Group VIB metal is calculated as oxide at 2-15%, preferably 5-10%, and the Group VIII metal is calculated as oxide at 2-10%, preferably 4-8%; the pore volume of the silicon-trapping catalyst is 0.5-1.0 mL / g, preferably 0.7-1.0 mL / g; the specific surface area is 250-500 m 2 / g, preferably 350-500m 2 / g; the acid content of B is 0.05-0.3mmol / g, preferably 0.08-0.2mmol / g; the preparation method of the catalyst is well known in the art, and one or more of the impregnation method, co-extrusion method, and co-precipitation method can be used. The hydrodeolefination catalyst is well known to those skilled in the art. As one of the specific embodiments, the active metal of the hydrodeolefination catalyst is a metal oxide of Group VIB and / or a metal oxide of Group VIII, the carrier is alumina or alumina modified with an additive, and based on the weight of the hydrodeolefination catalyst, the content of the metal oxide of Group VIB is 5%-30%, preferably 5%-15%, and the content of the metal oxide of Group VIII is 1%-15%, preferably 2%-6%; the specific surface area is 100-500m 2 / g, preferably 300-500m 2 / g, and a pore volume of 0.3-1.2 mL / g, preferably 0.4-0.8 mL / g. Methods for preparing the catalyst are well known in the art, and one or more of impregnation, co-extrusion, and co-precipitation methods may be employed. The hydrodeolefination catalyst may be a suitable commercial catalyst selected according to the requirements of the process flow, or may be prepared according to existing methods, or may be a regenerated catalyst obtained by regenerating a deactivated catalyst.

[0020] In the process of the present invention, the volume ratio of the catalysts in the hydrodesulfurization reaction zone and the desiliconization and deolefination reaction zone is 30:70 to 70:30, preferably 40:60 to 60:40.

[0021] In the process of the present invention, the hydrodesulfurization reaction zone and the desiliconization and deolefination reaction zone can be two beds in the same reactor, or can be arranged in different reactors.

[0022] In the process of the present invention, the organic sulfur content in the reaction product of the desiliconization and deolefination reaction zone is 50µg / g to 600µg / g.

[0023] In the process of the present invention, the gas separated from the gas-liquid is subjected to hydrogen sulfide removal and then mixed with new hydrogen and raw materials.

[0024] The process of the present invention has the following advantages compared with the prior art:

[0025] 1. In the existing technology for treating coker gasoline, hydrodesiliconization is generally performed before desulfurization. The purpose is to prevent silicon in the feed oil from contaminating the desulfurization catalyst and causing silicon poisoning of the desulfurization catalyst. The inventors broke with this convention. Based on the fact that silicon and olefins in coker gasoline are mainly concentrated in the light fraction, while the heavy fraction is mainly sulfur, they put the desulfurization reaction, which is usually set at the end, in front of the reaction, selectively allowing the heavy fraction to undergo the desulfurization reaction first. The depth of the desulfurization reaction is controlled to remove a portion of the sulfide in the heavy fraction. The desulfurized heavy fraction is then mixed with the light fraction to undergo deolefination and desiliconization reactions. This can not only reduce the strong exothermic heat of the large amount of olefins in the light fraction that would cause the reaction temperature to rise too high, but also avoid excessive desulfurization caused by the continuous temperature increase.

[0026] The coker gasoline hydrogenation process provided by this invention separates the feedstock into light and heavy fractions, which are then fed into different reaction zones. This effectively addresses the issue of poor quality coker gasoline, high levels of unsaturated hydrocarbons, and a relatively high silicon content, which can be susceptible to silicon poisoning during processing and thus impact the processing cycle. Furthermore, based on the requirement that the hydrodesiliconization reaction be optimized at a temperature above 275°C, the temperature in the desiliconization reaction zone is controlled to ensure optimal alignment of the silicon scavenger activity with the reactants.

[0027] 2. The heavy fraction provided by the present invention first enters the desulfurization reaction zone for hydrodesulfurization reaction, avoiding the strong exothermic phenomenon that occurs after the concentrated reaction of a large amount of olefins in the light fraction, which is beneficial to controlling the temperature rise and artificially reducing the difficulty of the reaction operation. At the same time, the desulfurization reaction zone product is mixed with the light fraction as a heat carrier and enters the desiliconization and deolefination reaction zone. The heavy fraction basically does not react in this reaction zone and only serves as a heat carrier to increase the absorption reaction heat capacity and reduce the effect of reaction temperature rise. At the same time, the desulfurization reaction zone effluent is mixed with the light fraction, and the light fraction, as a low-temperature material, can further reduce the reaction temperature and achieve reaction temperature matching. The process has a gentle reaction and strong controllability, and can better reduce the thermal cracking caused by excessive temperature, resulting in a decline in liquid yield and target product yield.

[0028] 3. The hydrodesulfurization reaction zone and the hydrodesiliconization and deolefination reaction zone provided by the present invention are respectively equipped with a hydrodesulfurization catalyst, a silicon scavenger, and a hydrodeolefination catalyst, and the reaction temperature is controlled to give full play to the desulfurization, deolefination activity, and desiliconization performance of the catalyst. This avoids the problem of the conventional coker gasoline hydrogenation process in which the initial silicon scavenger bed reaction temperature is too low, resulting in the silicon scavenger being unable to function and causing premature activity loss of the hydrogenation catalyst. It can also avoid passive over-desulfurization and reduce the sulfur injection amount of the downstream ethylene unit.

[0029] 4. This invention, guided by the desulfurization and desiliconization reactions, controls the reaction sequence of reactants and adopts a progressive desulfurization method. First, the difficult-to-remove sulfur is subjected to high-purity hydrogen to control the desulfurization depth. Then, the hydrogen sulfide concentration in the circulating hydrogen is controlled to suppress the desulfurization effect, avoid excessive desulfurization during the reaction, and effectively control the desulfurization depth. Furthermore, the desiliconization reaction temperature requirement is aligned with the control of the desulfurization reaction, achieving efficient matching of the reaction performance of the desulfurization catalyst, the silicon capture catalyst, and the deolefination catalyst.

[0030] 5. The present invention can be used to slightly modify existing coker gasoline hydrogenation or other similar hydrogenation equipment, which reduces the modification cost. At the same time, the operation steps are reduced, and the operation risk is also reduced. The coker gasoline hydrogenation process provided by the present invention has a simple process, convenient operation, good safety, environmental friendliness, and a long operating cycle. BRIEF DESCRIPTION OF THE DRAWINGS

[0031] Figure 1 It is a schematic diagram of the process flow of coking gasoline of the present invention. DETAILED DESCRIPTION

[0032] The method disclosed in the present invention is described in more detail below with reference to the accompanying drawings and specific embodiments. Figure 1 Many necessary equipment are omitted, such as oil pumps, valves, heating furnaces, etc.

[0033] like Figure 1 As shown, the heavy fraction of coker gasoline is mixed with the new hydrogen in pipeline 2 and the circulating hydrogen in pipeline 3 through pipeline 1, and enters the reactor from the top of the hydrotreating reactor 5. The material passes through the first reaction zone and undergoes hydrodesulfurization reaction. The light fraction of coker gasoline enters the middle part of the reactor through pipeline 4, mixes with the reaction effluent from the upper reaction zone, enters the second reaction zone, undergoes hydrodesiliconization and hydrodeolefination and degree of deolefination, and the mixed material flows out from the bottom of the reactor; enters the cold high fraction 7 through pipeline 6, and the liquid phase passes through pipeline 8 and enters the cold low fraction 9, and then enters the fractionation unit through pipeline 10; the circulating hydrogen separated in the cold high fraction 7 enters the circulating hydrogen desulfurization system through pipeline 11, and then passes through the circulating hydrogen compressor 12 and returns to pipeline 3.

[0034] The present invention will be further described below in conjunction with preferred embodiments. It should be understood that the preferred embodiments described herein are only used to illustrate and explain the present invention and are not used to limit the present invention.

[0035] The hydrorefining catalyst used in the first reaction zone in the Examples and Comparative Examples of the present invention was FH-40C, developed by the Fushun Petrochemical Research Institute and produced by the Fushun Branch of Sinopec Catalyst Company. The silicon scavenger used in the second reaction zone was FHRS-2, developed by the Fushun Petrochemical Research Institute and produced by the Fushun Branch of Sinopec Catalyst Company. The deolefination catalyst used in the second reaction zone was regenerated FH-40A, developed by the Fushun Petrochemical Research Institute and produced by the Fushun Branch of Sinopec Catalyst Company. The FH-40C catalyst was supported on alumina and contained W-Mo-Ni as the active metal components. The FHRS-2 catalyst was supported on expanded pore alumina and acid-modified with Mo-Ni as the active metal components. The FH-40A catalyst was supported on amorphous silica-alumina and acid-modified with Mo-Ni as the active metal components. ICP was used to analyze the composition content of the raw oil and the product, and nuclear magnetic resonance 29Si MAS NMR analysis showed that the silicon types in the raw oil and the product were siloxane compounds such as dimethylsiloxane, trimethylsiloxane and tetramethylsiloxane.

[0036] The conventional coker gasoline hydrofining process used in Comparative Example 1 of the present invention is a trickle bed reactor, where coker gasoline is mixed with hydrogen and enters the reactor for hydrodeolefination, desiliconization and desulfurization reactions. Figure 1 The hydrogenation process is a trickle bed reactor, but the first reaction zone is filled with silicon capture catalyst and hydrodeolefination catalyst in sequence, and the second reaction zone is filled with hydrodesulfurization catalyst. The light fraction is mixed with hydrogen and enters the first reaction zone for desiliconization and hydrodeolefination reaction. The first reaction product is mixed with the heavy fraction and enters the second reaction zone for hydrodesulfurization reaction. Figure 1 The hydrogenation process flow is shown in Table 1. The raw oil used is shown in Table 1, and the main operating process conditions and gasoline product properties are shown in Table 2.

[0037] Table 1 Properties of crude oil

[0038]

[0039] Table 2 Main operating process conditions and gasoline product properties at different times

[0040]

[0041] As can be seen from the data results of the Examples and Comparative Examples, the most significant feature of the present method is that it utilizes a conventional coker gasoline hydrorefining unit and a process grading method using a trickle-bed reactor. By using fraction cutting, the reaction fractions, reaction sequence, and various reaction types are deeply coupled, balancing reaction heat, reducing the effect of olefin exotherm on reaction temperature rise, controlling reaction depth, and maximizing the efficiency of the catalyst in each reaction zone. This reduces unit energy and hydrogen consumption, increases product liquid yield, reduces operational risks, and extends the operating cycle of the coker gasoline hydrorefining unit. This method brings considerable economic and social benefits to enterprises and possesses significant practical advantages.

Claims

1. A coker gasoline hydrogenation process, characterized in that The invention comprises the following contents: a coker gasoline feedstock is cut into a light fraction and a heavy fraction, the heavy fraction is mixed with hydrogen and enters a hydrodesulfurization reaction zone, contacts a hydrodesulfurization catalyst to carry out a hydrodesulfurization reaction, the hydrodesulfurization reaction effluent is mixed with the light fraction and enters a desiliconization and deolefination reaction zone, contacts a silicon scavenger and a hydrodeolefination catalyst to carry out desiliconization and deolefination reactions, the reaction product is subjected to gas-liquid separation, enters a fractionation zone, and is stripped and fractionated to obtain an ethylene cracking material; the hydrogen sulfide concentration in the hydrodesulfurization reaction effluent is 5000ppm to 20000ppm; the organic sulfur content in the hydrodesulfurization reaction effluent is 200µg / g to 1000µg / g; the volume ratio of the silicon scavenger to the hydrodeolefination catalyst in the desiliconization and deolefination reaction zone is 10:90 to 95:5; the hydrogen is fresh hydrogen and recycled hydrogen from the gas separated by gas-liquid separation after hydrogen sulfide is removed.

2. The hydrogenation process according to claim 1, wherein: The coker gasoline raw material has an initial boiling point of 30-100°C, a final boiling point of 160-260°C, an olefin content of 15wt%-50wt%, a silicon content of 3µg / g-500µg / g, and a sulfur content of 5000µg / g-15000µg / g.

3. The hydrogenation process according to claim 1 or 2, characterized in that: 40wt% to 100wt% of the coker gasoline raw material is coker gasoline.

4. The hydrogenation process according to claim 1, wherein: The cutting point temperature is 135-170° C., the sulfur content in the light fraction after cutting is 3000µg / g-8000µg / g, and the sulfur content in the heavy fraction is 7000µg / g-18000µg / g.

5. The hydrogenation process according to claim 1, wherein: The hydrodesulfurization reaction zone is a trickle bed reaction process, and the hydrodesulfurization reaction process conditions are as follows: inlet reaction temperature is 250℃~400℃, hydrogen partial pressure is 1.0MPa~16.0MPa, volume space velocity is 0.3h -1 ~16.0h -1 The volume ratio of hydrogen to oil is 100:1 to 2000:

1.

6. The hydrogenation process according to claim 1, wherein: The concentration of hydrogen sulfide in the circulating hydrogen in the hydrodesulfurization reaction zone is 0 ppm to 2000 ppm.

7. The hydrogenation process according to claim 1, wherein: The concentration of hydrogen sulfide in the hydrodesulfurization reaction effluent is 5000ppm to 15000ppm.

8. The hydrogenation process according to claim 1, wherein: The light fraction is mixed with the hydrodesulfurization reaction effluent and enters the desiliconization and deolefination reaction zone, or the light fraction is mixed with hydrogen and then mixed with the hydrodesulfurization reaction effluent and enters the desiliconization and deolefination reaction zone.

9. The hydrogenation process according to claim 1, wherein: The hydrodesulfurization catalyst uses a refractory porous oxide as a carrier. The hydrogenation active combination of the hydrodesulfurization catalyst is one or more of W, Mo, Ni and Co. The content of the hydrogenation active component is 3wt% to 55wt% based on the oxide. The shape is spherical or bar-shaped. The pore volume of the hydrodesulfurization catalyst is 0.4-1.0mL / g and the specific surface area is 200-400m 2 / g.

10. The hydrogenation process according to claim 1, wherein: The desiliconization and deolefination reaction zone is a trickle bed reaction process, and the reaction conditions are: inlet reaction temperature is 270℃~350℃, hydrogen partial pressure is 0.5MPa~16MPa, volume space velocity is 2h -1 ~6h -1 The volume ratio of hydrogen to oil is 100:1 to 2000:

1.

11. The hydrogenation process according to claim 1, wherein: The volume ratio of the silicon scavenger to the hydrogenation deolefination catalyst in the desiliconization and deolefination reaction zone is 35:65 to 80:

20.

12. The hydrogenation process according to claim 1, wherein: The active metals of the silicon scavenger are oxides of Group VIB metals and Group VIII metals. Based on the total weight of the silicon scavenger, the Group VIB metal is calculated as oxide at 2-15%, and the Group VIII metal is calculated as oxide at 2-10%. The pore volume of the silicon scavenger is 0.5-1.0 mL / g, and the specific surface area is 250-500 m 2 / g, the acid content of B is 0.05-0.3mmol / g.

13. The hydrogenation process according to claim 1, wherein: The active metal of the hydrodeolefination catalyst is a Group VIB metal oxide or / and a Group VIII metal oxide, the carrier is alumina or alumina modified with an additive, the Group VIB metal oxide content is 5%-30%, the Group VIII metal oxide content is 1%-15% based on the weight of the hydrodeolefination catalyst, and the specific surface area is 100-500m 2 / g, and the pore volume is 0.3-1.2mL / g.

14. The hydrogenation process according to claim 1, wherein: The volume ratio of the catalysts in the hydrodesulfurization reaction zone and the desiliconization and deolefination reaction zone is 30:70 to 70:

30.

15. The hydrogenation process according to claim 1, wherein: The hydrodesulfurization reaction zone and the desiliconization and deolefination reaction zone are two beds in the same reactor, or are arranged in different reactors.

16. The hydrogenation process according to claim 1, characterized in that: The organic sulfur content in the reaction product of the desiliconization and deolefination reaction zone is 50µg / g to 600µg / g.

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