Hydroprocessing process for coking gasoil and naphtha blend

Abstract

The application discloses a coking gasoline and diesel oil and naphtha mixed hydrogenation process method. The method comprises the following steps: (1) coking gasoline and diesel oil, naphtha and hydrogen are mixed into a diene pre-saturation reactor to perform olefin removal reaction; (2) the reaction material obtained in the step (1) is introduced into a gas phase reactor to perform desiliconization and hydrofining reaction; (3) the reaction material obtained in the step (2) is introduced into a liquid phase reactor after being pressurized, wherein the gas phase component is discharged upwards to obtain a light hydrogenation component, and the liquid phase component is discharged downwards to perform liquid phase hydrodearomatization reaction to obtain a heavy hydrogenation component; and (4) the light hydrogenation component and the heavy hydrogenation component obtained in the step (3) are mixed, subjected to a stripping and fractionation system, and diesel oil and naphtha products are obtained. The high-Si naphtha is mixed with the coking gasoline and diesel oil for processing, and through the setting of the gas phase reaction zone and the liquid phase reaction zone, the olefin hydrogenation reaction heat is fully utilized, the Si is efficiently removed, and the operation cycle of the coking gasoline and diesel oil processing device is prolonged.

Description

TECHNICAL FIELD

[0001] The present application belongs to the field of clean refining, and particularly relates to a coking gasoline and diesel oil and naphtha mixed hydrogenation process method. BACKGROUND

[0002] Delayed coking is a thermal cracking process that can convert heavy oil with high carbon residue into light oil. Its main products are coking naphtha, coking diesel, coking wax oil and coke. The coking oil has poor properties, mainly manifested in high olefin content and high nitrogen content, which will cause poor stability of the raw material, and a large amount of olefins will coking condense in the reaction process, seriously affecting the long-period stable operation of the device. The coking naphtha also contains silicon oil substances in the upstream silicon-containing defoaming agent, which will deposit on the surface of the catalyst during the hydrogenation reaction, causing permanent poisoning and deactivation of the catalyst.

[0003] Refineries generally use coking gasoline and diesel oil mixed hydrogenation, which can dilute the olefin content in the raw material, and also use the phase change process of high-boiling point substances in diesel components to absorb a large amount of latent heat of vaporization, which is beneficial to control the temperature rise of the catalyst bed and relatively prolong the operation cycle of the device. However, in the coking gasoline and diesel oil hydrogenation device, due to the high pressure, under the corresponding reaction conditions, part of the raw material in the reactor is vaporized, and the reaction is a gas-liquid-solid three-phase reaction, and the liquid film of heavy components covers the surface of the catalyst, hindering the adsorption reaction of light components on the surface of the catalyst, and the deposition of Si-containing species on the surface of the catalyst in the high-temperature zone has not been solved.

[0004] On the other hand, with the overcapacity of the fuel oil market, refining enterprises have begun to transform into chemical industry, especially to produce ethylene and aromatic basic chemical raw materials, and through hydrogenation technology, naphtha is converted into ethylene cracking raw materials and reforming feedstock to create more value. Due to the complex source of naphtha, the types of impurities are also very complex, especially the types of silicon compounds are various, among which the Si in coking naphtha is mainly in the form of cyclosiloxane, which is relatively easy to remove, while other naphtha may contain alkyl silane and silanol, which are difficult to remove, and the silicon content is as high as 300-1000 ppm. Current technology is difficult to remove such silicon to meet the requirements of reforming or ethylene feedstock, and new solutions need to be sought.

[0005] CN101591565A discloses a hydrofining method of poor quality gasoline. The gasoline raw oil is contacted with a hydrogenation protective agent under low temperature conditions, and the reaction effluent is mixed with the circulating oil and then contacted with a hydrodesilication agent and a hydrofining catalyst under high temperature conditions. The method can treat coking gasoline with high sulfur, high nitrogen and high olefin content, and the refined gasoline fraction can meet the feed requirements of the reforming pre-hydrogenation device and the steam cracking ethylene device. However, the method is not suitable for processing naphtha raw material with high silicon content, and the problem of catalyst bed coking caused by large heat release during the hydrogenation process of coking oil products has not been solved.

[0006] CN102051202A discloses a coking naphtha silicon capture agent and its application. The coking naphtha silicon capture agent uses alumina as a carrier, silicon dioxide as an additive, and W, Mo and Ni as hydrogenation components. The pore volume of the coking naphtha silicon capture agent is 0.5-0.70 mL / g, the specific surface area is 250-500 m 2 / g, the hydrogenation component content is 1%-20% in terms of oxide, and the acid content is 0.3-0.5 mmol / g. The coking naphtha silicon capture agent can effectively remove impurities such as silicon in coking naphtha, and protect the coking naphtha hydrofining catalyst from permanent deactivation caused by silicon poisoning. However, the silicon capture agent is mainly aimed at removing siloxane silicon compounds in coking distillate oil, and it is difficult to remove siloxane and silane silicon compounds. SUMMARY

[0007] In view of the short running cycle of coking gasoline and diesel oil hydrogenation and the high Si content in some naphtha sources, which is difficult to remove, the present application provides a coking gasoline and diesel oil and naphtha mixed hydrogenation process method. High Si naphtha is mixed with coking gasoline and diesel oil for processing, and through the setting of gas phase reaction zone and liquid phase reaction zone, the hydrogenation reaction heat of olefins is fully utilized to achieve efficient removal of Si, and the running cycle of the coking gasoline and diesel oil processing device is prolonged.

[0008] The present application provides a coking gasoline and diesel oil and naphtha mixed hydrogenation process method, which comprises the following steps:

[0009] (1) coking gasoline, diesel oil, naphtha and hydrogen gas are mixed into a diene pre-saturation reactor to perform a de-olefin reaction;

[0010] (2) the reaction material obtained in step (1) is introduced into a gas phase reactor to perform a desiliconization and hydrofining reaction;

[0011] (3) the reaction material obtained in step (2) is pressurized and introduced into a liquid phase reactor, wherein the gas phase component is discharged upward and introduced into a high pressure separator to obtain a hydrogenated light component, and the liquid phase component is discharged downward to perform a liquid phase hydrodearomatization reaction to obtain a hydrogenated heavy component which is discharged from the bottom of the liquid phase reactor;

[0012] (4) The hydrogenated light component obtained in step (3) is mixed with a hydrogenated heavy component, and then subjected to a stripping and fractionation system to obtain diesel and naphtha products.

[0013] Further, the proportion of coking gasoline in the coking gasoline-diesel is 20wt% to 90wt%, preferably 30wt% to 70wt%, and the proportion of coking diesel in the coking gasoline-diesel is 10wt% to 80wt%, preferably 30wt% to 70wt%. The coking gasoline-diesel has a sulfur content of ≯15000 μg / g, preferably 6000 to 12000 μg / g, a nitrogen content of ≯400 μg / g, preferably 100 to 380 μg / g, and a polycyclic aromatic hydrocarbon content of ≯50wt%, preferably 10wt% to 35wt%.

[0014] Further, the naphtha is high-Si naphtha, which can be at least one of straight-run naphtha, coking naphtha, and hydrogenation modified naphtha. The high-Si naphtha contains at least one silicon-containing compound other than cyclic siloxane, such as one or more of tetramethylsilane, dimethoxydimethylsilane, and tetraethylsilane. The total Si content in the high-Si naphtha is 1 to 3000 μg / g. The total Si content in the silicon-containing compound other than cyclic siloxane is 1 to 2800 μg / g, preferably 300 to 1000 μg / g. The high-Si naphtha has a sulfur content of ≯1000 μg / g, preferably 200 to 600 μg / g, and a nitrogen content of ≯200 μg / g, preferably 10 to 150 μg / g.

[0015] Further, the mass ratio of the coking gasoline-diesel to the naphtha is 1 to 5:1.

[0016] Further, at the initial stage of the operation, the coking gasoline-diesel, the naphtha, and hydrogen are heated by a heating furnace and then fed into the diene pre-saturation reactor. When the temperature of the gas phase reactor increases significantly and the outlet temperature of the gas phase reactor reaches 100°C to 140°C, the heating furnace is stopped and the feed (the coking gasoline-diesel, the naphtha, and hydrogen) and the reaction material obtained at the outlet of the gas phase reactor are heated by a heat exchanger to meet the feed temperature requirement of the diene pre-saturation reactor.

[0017] Further, the reaction conditions of the diene pre-saturation reactor include a reaction pressure of 0.1 to 4.0 MPa, preferably 0.5 to 3.0 MPa; a hydrogen to oil volume ratio of 100:1 to 600:1, preferably 150:1 to 500:1; a volume space velocity of 1.0 to 10.0 h -1 , preferably 2.0 to 8.0 h -1 ; and a reaction temperature of 50 to 200°C, preferably 80 to 160°C.

[0018] Further, the catalyst loaded in the diene pre-saturation reactor is a light distillate oil hydrogenation catalyst, such as the FH-40A and FH-40B catalysts developed by FRIPP.

[0019] Further, the reaction conditions of the gas phase reactor include: a reaction pressure of 0.5-4.0 MPa, preferably 1.0-3.0 MPa; a hydrogen to oil volume ratio of 100:1-1000:1, preferably 150:1-450:1; a volume space velocity of 0.1-10.0 h -1 , preferably 1.0-6.0 h -1 ; and a reaction temperature of 150-350°C, preferably 220-320°C.

[0020] Further, the gas phase reactor is loaded with a silicon capture catalyst and a light distillate oil hydrogenation catalyst. The silicon capture catalyst is loaded first, followed by the light distillate oil hydrogenation catalyst, in the direction of material flow. The volume ratio of the silicon capture catalyst to the light distillate oil hydrogenation catalyst is 1:5-3:1.

[0021] Further, the silicon capture catalyst is well known to those skilled in the art, and has a VIB group metal or / and a VIII group metal as an active metal, the VIB group metal is preferably molybdenum and / or tungsten, and the VIII group metal is preferably nickel and / or cobalt, and an alumina or an alumina modified with an additive as a carrier. The content of the VIB group metal in the form of an oxide is 5%-30%, preferably 5%-15%, and the content of the VIII group metal in the form of an oxide is 1%-15%, preferably 2%-6%, based on the weight of the silicon capture catalyst. The specific surface area is 100-500 m 2 / g, preferably 300-500 m 2 / g, and the pore volume is 0.3-1.2 mL / g, preferably 0.4-0.8 mL / g, such as the FHRS-2 silicon capture catalyst developed by FRIPP. The light distillate oil hydrogenation catalyst is well known to those skilled in the art, and has a VIB group metal or / and a VIII group metal as an active metal, and an alumina or an alumina modified with an additive as a carrier. The content of the VIB group metal in the form of an oxide is 5%-30%, preferably 10%-20%, and the content of the VIII group metal in the form of an oxide is 1%-10%, preferably 2%-8%, based on the weight of the light distillate oil hydrogenation catalyst. The specific surface area is 100-400 m 2 / g, preferably 200-300 m 2 / g, and the pore volume is 0.2-1.0 mL / g, preferably 0.4-0.8 mL / g, such as the FH-40A and FH-40B catalysts developed by FRIPP.

[0022] Further, the light distillate oil hydrogenation catalysts filled in the diene olefin pre-saturation reactor and the light distillate oil hydrogenation catalysts filled in the gas phase reactor can be the same or different.

[0023] Further, the gas phase reactor can adopt two or more reactors in parallel and operate in rotation, when the catalyst activity in one of the gas phase reactors is significantly reduced, switching to another parallel gas phase reactor to continue the reaction. In the gas phase reactor, mainly the saturation of mono-olefins, the removal of Si compounds in coked gasoline and diesel, and the removal of Si compounds in high-Si naphtha, as well as the removal of small molecular impurities (S, N) in naphtha fraction, etc. occur.

[0024] Further, a compressor is provided between the gas phase reactor and the liquid phase reactor for pressurization, which can be a conventional commercial compressor, such as a reciprocating compressor or a centrifugal compressor. Among them, the pressurization can only ensure the normal feeding of the liquid phase reactor and meet the operating pressure requirements of the liquid phase reactor.

[0025] Further, a flash zone is provided in the liquid phase reactor, which is not filled with catalyst in the flash zone and above, and is a reaction zone below. The reaction material obtained from the gas phase reactor is fed to the flash zone of the liquid phase reactor after being pressurized, the gas phase components are discharged upward from the liquid phase reactor, and the liquid phase components are subjected to isomerization and hydrodearomatization reaction downward, and the heavy hydrogenated components are discharged from the bottom of the liquid phase reactor.

[0026] Further, the main reaction in the liquid phase reactor is deep de-aromatization. The liquid phase reactor is filled with diesel hydrogenation catalyst, which is well known to those skilled in the art. The active metal of the diesel hydrogenation catalyst is a Group VIB metal or / and a Group VIII metal, the Group VIB metal is preferably molybdenum and / or tungsten, and the Group VIII metal is preferably nickel and / or cobalt. The content of the Group VIB metal, calculated as the oxide, is 5% to 30%, preferably 15% to 25%, and the content of the Group VIII metal, calculated as the oxide, is 1% to 15%, preferably 3% to 8%, based on the weight of the diesel hydrogenation catalyst; the carrier is alumina or alumina modified with an additive, the additive is one or more of B, P, Mg, Zr or Si, and the content of the additive, calculated as the oxide, is 3% to 15%, preferably 3% to 10%, based on the weight of the carrier, such as the FHUDS-7 and FHUDS-10 catalysts developed by FRIPP.

[0027] Further, the reaction conditions of the liquid phase reactor include: the reaction pressure is 4.0 to 7.5 MPa, preferably 5.0 to 7.0 MPa; the volume space velocity is 0.1 to 5.0 h -1 , preferably 0.5 to 3.0 h -1The reaction temperature is 280-400 DEG C, preferably 300-360 DEG C. The reaction pressure of the liquid phase reactor is at least 1-8 MPa, preferably 2-6 MPa, higher than that of the gas phase reactor.

[0028] Compared with the prior art, the present application has the following advantages:

[0029] The process method of the present application aims at the problem of long-period operation limitation of a coking gasoline and diesel oil device, and designs a low-energy-consumption and low-cost process line, fully utilizes reaction heat, and can process high-Si naphtha inferior oil, and provides qualified chemical raw materials and vehicle diesel oil products.

[0030] The present application uses coking gasoline and diesel oil and high-Si naphtha as mixed raw oil. Since the Si species in the high-Si naphtha is enriched in the light component, it overlaps with the olefins in the coking gasoline and diesel oil in the distillation range, and needs a higher reaction temperature for removal, so the reaction heat of the olefin hydrogenation in the coking oil processing process can be fully utilized, and higher energy consumption caused by separate processing of high-Si oil is avoided, which is beneficial to joint processing.

[0031] The present application only needs to use a heating furnace to preheat the raw oil at the initial stage of start-up. Since the olefin content in the coking gasoline and diesel oil is high, the rapid reaction heat release can increase the outlet stream temperature of the reactor, at this time, the heating furnace can be stopped, and the heat exchange between the normal temperature raw material and the gas reactor outlet material can meet the feeding temperature, thereby fully utilizing the reaction heat to save energy and reduce consumption.

[0032] The method of the present application controls the reaction conditions of the gas phase reactor, not only for gas phase reaction, but more for better cooperation with the liquid phase reactor, so as to be more conducive to the reaction environment of subsequent diesel oil deep desulfurization and de-aromatics. Compared with the conventional fixed bed hydrogenation technology, the method of the present application can omit the heat exchange and compressor setting in the hydrogen circulation process, significantly reducing energy consumption and construction investment. Compared with the liquid phase hydrogenation technology, the present application utilizes the oil and hydrogen in the gas phase state at the outlet of the gas phase reactor, high-temperature pressurized liquefaction, so that hydrogen is more preferentially dissolved in the oil, and the setting of the hydrogen mixer can be omitted, reducing the construction investment. At the same time, the energy consumption of the whole reaction system of the present application is low, and the safety is improved, and the requirement for the reactor material can be reduced. From the perspective of chemical reaction rate, the present application is a combination of gas phase and liquid phase reactions, and the reaction rate is significantly improved. BRIEF DESCRIPTION OF DRAWINGS

[0033] Figure 1 A schematic diagram of the gasoline and diesel oil mixed hydrogenation process used in examples 1-3 is shown in the figure.

[0034] Among them, coking gasoline and diesel, naphtha and hydrogen-1; heat exchanger-2; diene pre-saturated reactor-3; gas phase reactor-4, 5; material obtained from gas phase reactor-6; compressor-7; liquid phase reactor-8; hydrogenated heavy component-9; gas phase component-10; high pressure separator-11; hydrogenated light component-12; hydrogen containing hydrogen sulfide-13; steam top fractionation system-14. Detailed Implementation

[0035] The present invention will be further described below with reference to the embodiments, but it should be understood that the scope of protection of the present invention is not limited to the embodiments.

[0036] In this invention, unless otherwise explicitly stated, percentages and whole contents are all expressed by mass. The following is in conjunction with... Figure 1 The process flow of the present invention will be described in detail.

[0037] Coking gasoline and diesel, naphtha, and hydrogen 1 are heated in a heater (not shown in the figure) and then enter a diene pre-saturated reactor 3 under certain temperature and pressure conditions. After that, they enter a gas phase reactor 4 or 5 operating in parallel. The material 6 obtained from the gas phase reactor passes through a heat exchanger 2 and is then pressurized by a compressor 7 before entering a liquid phase reactor 8. The liquefied heavy components enter the reaction zone to continue reacting and obtain hydrogenated heavy components 9. The gas phase components 10 enter a high-pressure separator 11, where they are separated into hydrogenated light components 12 and hydrogen containing hydrogen sulfide 13. The hydrogenated heavy components 9 and hydrogenated light components 12 are mixed and enter a vapor strip and fractionation system 14 to finally obtain diesel products and naphtha as chemical raw materials.

[0038] Examples 1-3

[0039] Adopting such Figure 1 The process flow diagram shows that a catalyst bed is set up in the diene presaturation reactor, filled with 20 mL of light distillate oil hydrogenation catalyst; a catalyst bed is set up in the gas phase reactor, filled with 60 mL of silica-collecting catalyst B and light distillate oil hydrogenation catalyst A in the direction of material flow, with a volume ratio of catalyst B to catalyst A of 1:2; a catalyst bed is set up in the liquid phase reactor, filled with 60 mL of diesel hydrogenation catalyst C. The feedstock is a mixture of high-Si naphtha, coking gasoline, and coking diesel oil, with a mass ratio of 20:50:30. The high-Si naphtha contains tetramethylsilane, dimethoxydimethylsilane, and tetraethylsilane, with a total Si content of 306 μg / g. The properties of the feedstock are shown in Table 1, and the properties of the catalyst are shown in Table 2.

[0040] Comparative Example 1

[0041] The process flow of gas phase hydrogenation is adopted, and a diene pre-saturation reactor and a gas phase hydrogenation reactor are connected in series. The diene pre-saturation reactor is filled with 20 mL of light distillate oil hydrogenation catalyst A, and the gas phase hydrogenation reactor is filled with 20 mL of silicon capture catalyst B, 40 mL of light distillate oil hydrogenation catalyst A and 60 mL of diesel oil hydrogenation catalyst C in a gradient manner. A heat exchanger and a condenser are arranged after the gas phase reactor, and the unliquefied gas phase can be recycled after desulfurization. The liquefied components are introduced into the stripping and fractionation equipment again to obtain naphtha and diesel products. The properties of the raw oil and the catalysts are the same as those in Examples 1-3, and the reaction process conditions and results are shown in Table 3.

[0042] Comparative Example 2

[0043] The conventional fixed bed three-phase hydrogenation reaction process is adopted, and a diene pre-saturation reactor and a fixed bed reactor are connected in series. The diene pre-saturation reactor is filled with 20 mL of light distillate oil hydrogenation catalyst A, and the fixed bed reactor is filled with 20 mL of silicon capture catalyst B, 40 mL of light distillate oil hydrogenation catalyst A and 60 mL of diesel oil hydrogenation catalyst C in a gradient manner. After each reactor, high and low separation, stripping and other devices are arranged according to the conventional method. Hydrogen is pressurized by a circulating hydrogen compressor after removing hydrogen sulfide and recycled. The properties of the raw oil and the catalysts are the same as those in Examples 1-3, and the reaction process conditions and results are shown in Table 3.

[0044] Comparative Example 3

[0045] The hydrogenation process flow of Example 1 is adopted, and the only difference is that a high-power reciprocating compressor is arranged between the gas phase and liquid phase reactors, and the gas phase reactor effluent is completely liquefied into the liquid phase reactor. The reaction process conditions and results are shown in Table 3.

[0046] Table 1 Properties of raw oil

[0047] Oil properties High Si naphtha Coker gasoil Distillation range, °C 45～180 55～361 [S, pg g -1 ]] 337 10061 [N, pg g -1 ]]> 45 365 Total Si, pg-g -1 ]] 437 12 Polycyclic aromatic content, wt% - 15

[0048] Table 2 Catalyst properties

[0049] Catalyst type A B C Grade FH-40A FHRS-2 FHUDS-10 Active metal Mo-Ni Mo-Ni Mo-Ni Diameter, mm 2 2 1.2 Specific surface area, m 2 ·g -1 ]]> 360 220 230 Pore volume, mL-g -1 ]] 0.53 0.44 0.33

[0050] Table 3 Process conditions and results

[0051]

[0052]

[0053] The evaluation results in Table 3 can show that the processing method of the application can realize efficient Si removal by matching and combining gas phase and liquid phase reactions, and using high Si naphtha and coking gasoline and diesel oil for joint processing, and can prolong the operation cycle of the coking oil processing device. After ten months of operation, the Si content of the product is still low, indicating that the catalyst has not been poisoned by Si. The Si content of the product of the comparative example increases significantly after ten months, indicating that the catalyst activity has been attenuated and it is difficult to achieve long-period operation. If only the gas phase hydrogenation technology is used, it is not suitable for hydrodesulfurization and de-aromatics of diesel oil, and the quality of the oil product cannot meet the standard; using conventional gas-liquid-solid three-phase hydrogenation technology, Si starts to deposit on the surface of the main catalyst in large quantities only in the high temperature zone at the lower part of the reactor, affecting the product quality and the Si removal effect is poor; in Comparative Example 3, a high-power compressor is used to liquefy all the gas phase reactor effluent into the liquid phase reactor, which increases the liquid phase reactor space velocity, reduces the reaction effect, and significantly increases the energy consumption.

Claims

1. A method for hydrogenating a mixture of coking gasoline / diesel and naphtha, comprising the following steps: (1) Coking gasoline, diesel, naphtha and hydrogen are mixed and fed into a diene presaturated reactor for deolefination reaction; (2) The reactants obtained in step (1) enter the gas phase reactor for desilication and hydrogenation purification reactions; (3) The reactants obtained in step (2) are pressurized and then enter the liquid phase reactor. The gas phase component is discharged upward and enters the high pressure separator to obtain the hydrogenated light component. The liquid phase component is discharged downward to undergo liquid phase hydrogenation and dearomatization reaction to obtain the hydrogenated heavy component, which is discharged from the bottom of the liquid phase reactor. (4) The hydrogenated light component obtained in step (3) is mixed with the hydrogenated heavy component and then passed through a vaporizer to a fractionation system to obtain diesel and naphtha products; Based on the mass of the coking gasoline and diesel, the proportion of coking gasoline in the coking gasoline and diesel is 20wt%~90wt%, and the proportion of coking diesel in the coking gasoline and diesel is 10wt%~80wt%. The naphtha is a high-Si naphtha, which includes at least one non-cyclic siloxane silicon-containing compound. The reaction conditions of the gas-phase reactor include: reaction pressure 0.5~4.0 MPa, hydrogen-to-oil volume ratio 100:1~1000:1, and volume hourly space velocity 0.1~10.0 h⁻¹. -1 The reaction temperature is 150~350℃; The gas-phase reactor is filled with a silicon-scavenging catalyst and a light distillate oil hydrogenation catalyst; the silicon-scavenging agent and the light distillate oil hydrogenation catalyst are filled sequentially in the direction of material flow. The reaction conditions of the liquid-phase reactor include: reaction pressure 4.0~7.5MPa, and volumetric hourly space velocity 0.1~5.0h⁻¹. -1 The reaction temperature is 280~400℃.

2. The method according to claim 1, characterized in that, Based on the mass of the coking gasoline and diesel, the proportion of coking gasoline in coking gasoline and diesel is 30wt%~70wt%, and the proportion of coking diesel in coking gasoline and diesel is 30wt%~70wt%.

3. The method according to claim 1 or 2, characterized in that, The sulfur content of the coking gasoline and diesel is ≤15000μg / g, the nitrogen content is ≤400μg / g, and the polycyclic aromatic hydrocarbon content is ≤50wt%.

4. The method according to claim 3, characterized in that, The coking gasoline and diesel have a sulfur content of 6000~12000μg / g, a nitrogen content of 100~380μg / g, and a polycyclic aromatic hydrocarbon content of 10wt%~35wt%.

5. The method according to claim 1, characterized in that, The silicon-containing compounds of the noncyclic siloxane class are one or more of tetramethylsilane, dimethoxydimethylsilane, and tetraethylsilane.

6. The method according to claim 1 or 5, characterized in that, The total Si content in the silicon-containing compounds of noncyclic siloxanes is 1~2800 μg / g, and the sulfur content of the high-Si naphtha is ≤1000 μg / g and the nitrogen content is ≤200 μg / g.

7. The method according to claim 6, characterized in that, The total Si content in the silicon-containing compounds of noncyclic siloxanes is 300~1000 μg / g, and the sulfur content of the high-Si naphtha is 200~600 μg / g, and the nitrogen content is 10~150 μg / g.

8. The method according to claim 1, characterized in that, The mass ratio of coking gasoline / diesel to naphtha is 1~5:

1.

9. The method according to claim 1, characterized in that, The reaction conditions of the diene presaturated reactor include: reaction pressure 0.1~4.0 MPa, hydrogen-to-oil volume ratio 100:1~600:1, and volume hourly space velocity 1.0~10.0 h⁻¹. -1 The reaction temperature is 50~200℃.

10. The method according to claim 9, characterized in that, The reaction conditions of the diene presaturated reactor include: a reaction pressure of 0.5–3.0 MPa, a hydrogen-to-oil volume ratio of 150:1–500:1, and a volume hourly space velocity of 2.0–8.0 h⁻¹. -1 The reaction temperature is 80~160℃.

11. The method according to claim 1, characterized in that, The reaction conditions of the gas-phase reactor include: reaction pressure 1.0~3.0 MPa, hydrogen-to-oil volume ratio 150:1~450:1, and volume hourly space velocity 1.0~6.0 h⁻¹. -1 The reaction temperature is 220~320℃.

12. The method according to claim 1, characterized in that, The volume ratio of the silicon-capturing catalyst to the light distillate oil hydrogenation catalyst is 1:5 to 3:

1.

13. The method according to claim 1, characterized in that, The reaction conditions of the liquid-phase reactor include: reaction pressure 5.0~7.0 MPa, and volumetric hourly space velocity 0.5~3.0 h⁻¹. -1 The reaction temperature is 300~360℃.

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