Processing method of delayed coking light naphtha

Abstract

The invention discloses a delayed coking light naphtha processing method. The method comprises the following steps: (I) delayed coking light naphtha and hydrogen are mixed and then enter a hydrofining reaction zone for a hydrofining reaction, a refined naphtha material flow is obtained, and the hydrofining zone is sequentially filled with a silicon capturing agent and a hydrofining catalyst in the material flowing direction; (II) the refined naphtha material flow obtained in the step (I) enters a hydro-conversion reaction zone for hydro-conversion reaction, and converted delayed coking light naphtha is obtained. According to the method disclosed by the invention, the saturation and conversion of the delayed coking light naphtha are realized, a hydrocarbon stream product which does not contain olefin and has improved n-alkane content is obtained, and the yield of triene is higher when the hydrocarbon stream product is used as an ethylene raw material.

Description

Technical Field

[0001] This invention relates to the field of petrochemical technology, specifically to a delayed coking method for processing light naphtha (mainly referring to hydrocarbons rich in C5 and C6 hydrocarbons). Background Technology

[0002] Delayed coking technology offers good economic benefits due to its lower investment cost and the fact that the coking products, after hydrorefining, are excellent oil products and feedstocks for secondary processing. Products from delayed coking units can be categorized by distillation range into rich coking gas, coking naphtha, coking diesel, coking distillate oil, and petroleum coke.

[0003] Coking naphtha contains some light naphtha with C5 and C6 as the main components, which is usually used as a gasoline fraction. However, due to its low distillation range, it affects the saturated vapor pressure of gasoline products. At the same time, with the continuous growth of ethylene demand, using delayed coking light naphtha as ethylene feedstock meets the requirements of "oil conversion". However, there are two problems when using delayed coking light naphtha as ethylene feedstock: (1) Delayed coking light naphtha has a high olefin content (bromine value greater than 50gBr2 / 100g), which is easy to coke. After entering the steam cracking ethylene unit, it will shorten the coking cycle and affect the economic benefits of the ethylene unit; (2) Delayed coking light naphtha contains a certain amount of isomeric hydrocarbons (about 20wt% to 40wt%), which will reduce the yield of trienes in the ethylene unit and reduce its economic benefits. In addition, due to the high silicon content of delayed coking light naphtha, it poses challenges to subsequent refining and conversion processes.

[0004] CN201680075145.5 provides a method and system for producing olefins and aromatics. The method may include removing silica from a coking naphtha feedstock to produce a first effluent, hydrogenating the first effluent to produce a second effluent, reacting the second effluent to produce a third effluent containing aromatics, a fourth effluent containing olefins, and a fifth effluent, separating the fourth effluent to produce a propylene product stream, an ethylene product stream, and a sixth effluent, and recycling the sixth effluent by combining it with the second effluent. Summary of the Invention

[0005] To address the shortcomings of existing technologies, this invention provides a delayed coking light naphtha processing method. The product obtained by processing delayed coking light naphtha using this method has fully saturated olefins and increased n-hydrocarbon content, which can significantly improve the yield of trienes (ethylene, propylene, and butadiene) when used as a feedstock for ethylene production by steam cracking.

[0006] This invention provides a method for processing delayed coking light naphtha, the method comprising:

[0007] (I) Delayed coking light naphtha is mixed with hydrogen and then enters the hydrorefining reaction zone for hydrorefining reaction to obtain refined naphtha stream. The hydrorefining zone is sequentially filled with silica scavenger and hydrorefining catalyst along the material flow direction.

[0008] (II) The refined naphtha stream from step (I) enters the hydroconversion reaction zone for hydroconversion reaction to obtain the converted delayed coking light naphtha.

[0009] Further, the properties of the delayed coking light naphtha are as follows: initial boiling point is 15–30°C, and final boiling point is 60–90°C. Further, the silicon content (SiO2) in the delayed coking light naphtha is 5–20 μg / g, and the bromine value is 30–70 gBr / 100g. Further, the C5-C6 isomeric hydrocarbon content in the delayed coking light naphtha is 20 wt%–50 wt%, preferably 30 wt%–50 wt%, based on a total weight of 100 wt% of the delayed coking light naphtha.

[0010] Further, in step (I), the bromine value of the refined naphtha stream is 0.08 to 0.30 gBr / 100g, preferably 0.15 to 0.30 gBr / 100g, for example, 0.15 gBr / 100g, 0.16 gBr / 100g, 0.17 gBr / 100g, 0.18 gBr / 100g, 0.20 gBr / 100g, 0.22 gBr / 100g, 0.24 gBr / 100g, 0.26 gBr / 100g, 0.28 gBr / 100g, 0.30 gBr / 100g, and any range between any two values.

[0011] Further, in step (I), the silicon-collecting agent can be a conventional silicon-collecting agent in the art, generally including alumina and a hydrogenated metal, wherein the hydrogenated metal is selected from Group VIII and Group VIB, with Group VIB metals preferably being molybdenum and / or tungsten, and Group VIII metals preferably being cobalt and / or nickel. The silicon-collecting agent can be, for example, the FHRS series silicon-collecting agent developed by Sinopec (Dalian) Petrochemical Research Institute Co., Ltd., such as FHRS-3.

[0012] Further, in step (I), the hydrorefining catalyst can be a conventional hydrorefining agent in the art, generally comprising a hydrorefining active metal component and a support. The hydrorefining active metal is a Group VIB metal or a Group VIII metal, preferably molybdenum and / or tungsten for Group VIB, and preferably cobalt and / or nickel for Group VIII. The support is an alumina support. Generally, based on the weight of the catalyst, the content of the Group VIB metal (calculated as oxide) is 20%–30%, and the content of the Group VIII metal (calculated as oxide) is 2%–10%. The hydrorefining catalyst can be an FF series catalyst developed by Sinopec (Dalian) Petrochemical Research Institute Co., Ltd., such as FF-66.

[0013] Furthermore, in step (I), the volume ratio of the silicon-scavenging agent to the hydrorefining catalyst is 1:4 to 1:10.

[0014] Further, in step (I), the hydrorefining conditions are: reaction temperature 100–250°C, reaction pressure 2–8 MPa, and liquid hourly space velocity (LHSV) 8.0–15.0 h⁻¹. -1 The hydrogen-to-oil volume ratio is 100:1 to 2000:1.

[0015] Further, in step (II), the hydrogenation conversion conditions are: reaction temperature 300–450°C, reaction pressure 2–8 MPa, and liquid hourly space velocity (LHSV) 0.1–5.0 h⁻¹. -1 The hydrogen-to-oil volume ratio is 100:1 to 2000:1.

[0016] Further, in step (II), the hydroconversion reaction zone is filled with a hydroconversion catalyst. Further, the volume ratio of the hydroconversion catalyst to the hydrorefining catalyst is 3:1 to 10:1.

[0017] Further, the hydroconversion catalyst comprises modified ZSM-5 molecular sieve and macroporous alumina supported on Group VIB and Group VIII metals. Further, based on the weight of the hydroconversion catalyst, the content of modified ZSM-5 molecular sieve is 35 wt% to 80 wt%, the content of macroporous alumina is 15 wt% to 50 wt%, and the total content of Group VIB and Group VIII metals (based on oxides) is 0.5 wt% to 8.0 wt%. Preferably, based on the weight of the hydroconversion catalyst, the content of modified ZSM-5 molecular sieve is 50 wt% to 75 wt%, the content of macroporous alumina is 20 wt% to 30 wt%, and the total content of Group VIB and Group VIII metals (based on oxides) is 1.0 wt% to 5.0 wt%. Further, the Group VIB metal is preferably molybdenum and / or tungsten, and the Group VIII metal is preferably cobalt and / or nickel. Furthermore, in the macroporous alumina loaded with Group VIB and Group VIII metals, the mass ratio of the Group VIB metal (calculated as oxide) to the Group VIII metal (calculated as oxide) is 2.0 to 6.0.

[0018] Furthermore, the mesoporous pore volume of the modified ZSM-5 molecular sieve accounts for 50% to 80% of the total pore volume, preferably 60% to 70%.

[0019] Furthermore, the total content C of Group VIB and Group VIII metals on the surface of the hydroconversion catalyst is... ms The total content of Group VIB and Group VIII metals in the catalyst bulk phase, C mt The molar ratio is 0.60–0.90. The total amount of Group VIB and Group VIII metals on the catalyst surface, C ms This refers to the total molar amount, in atomic terms, of Group VIB and Group VIII metals per gram of catalyst surface. The total amount of Group VIB and Group VIII metals in the catalyst bulk phase, C. mt It refers to the total molar amount of Group VIB and Group VIII metals per gram of catalyst bulk phase, expressed in atomic quantities.

[0020] Furthermore, the preparation method of the hydrogenation conversion catalyst includes the following steps:

[0021] (1) Pore expansion treatment of ZSM-5 molecular sieve;

[0022] (2) Remove non-framework aluminum from step (1) to obtain modified ZSM-5 molecular sieve;

[0023] (3) Macroporous alumina was impregnated with an impregnation solution containing Group VIB and Group VIII metals and then calcined to obtain macroporous alumina loaded with Group VIB and Group VIII metals.

[0024] (4) The modified ZSM-5 molecular sieve and macroporous alumina loaded with Group VIB and Group VIII metals were mixed, shaped, and calcined to obtain a hydroconversion catalyst.

[0025] Further, in step (1), the ZSM-5 molecular sieve can be a commercially available product or a microporous hydrogen-form ZSM-5 molecular sieve prepared according to existing technology. The properties of the ZSM-5 molecular sieve are as follows: SiO2 / Al2O3 molar ratio 30-100, specific surface area 300-450 m² / g. 2 / g, pore volume 0.15~0.20cm³ 3 / g.

[0026] Further, in step (1), the pore-expanding treatment can be one or more of acid treatment, alkali treatment, and hydrothermal treatment, preferably hydrothermal treatment. The temperature of the hydrothermal treatment is 400-700℃, preferably 500-600℃, the time is 0.5-5h, preferably 1-2h, and the pressure is 0.05-0.5MPa, preferably 0.1-0.3MPa.

[0027] Further, in step (2), the method for removing non-skeletal aluminum can be a buffer solution method. The buffer solution used is one or more of oxalic acid-ammonium oxalate solution and acetic acid-ammonium acetate solution. The pH value of the buffer solution is 4.5-6.5, preferably 5.0-6.0. The molar concentration of the organic acid in the buffer solution is 0.1-1.0 mol / L. The liquid-solid volume ratio of the buffer solution to the molecular sieve obtained in step (1) is 3:1-10:1.

[0028] Further, in step (2), the specific process for removing non-skeletal aluminum is as follows: the molecular sieve obtained in step (1) is mixed with a buffer solution and stirred, the treatment temperature is 40-80℃, the treatment time is 0.5-3h, and then solid-liquid separation (e.g., vacuum filtration) is performed; and the above operation is repeated 2-4 times.

[0029] Furthermore, the impregnation in step (3) is performed using a pressure impregnation method.

[0030] Furthermore, the pressure impregnation includes: mixing macroporous alumina with an impregnation solution containing Group VIB and Group VIII metals, and impregnating in a closed atmosphere under nitrogen.

[0031] Furthermore, the impregnation conditions include: a pressure of 0.5–3.0 MPa, a temperature of 60–80 °C, and an impregnation time of 1–4 h.

[0032] Further, the impregnation is preferably carried out under stirring at a stirring rate of 100–300 rpm. Further, after impregnation, stirring and heating are stopped, and the mixture is allowed to stand for 2–4 hours until cooled to room temperature. The resulting material is first subjected to conventional filtration and drying, and then calcined to obtain macroporous alumina loaded with Group VIB and Group VIII metals. The drying temperature is 100°C–150°C, and the drying time is 2–4 hours; the calcination temperature is 400°C–600°C, and the calcination time is 3–5 hours.

[0033] Furthermore, in step (3), the pore volume of the macroporous alumina is 0.8–1.5 cm³. 3 / g, specific surface area of ​​200-400m² 2 / g.

[0034] Further, in step (3), the impregnation solution containing Group VIB and Group VIII metals contains one or more of tungsten or molybdenum salts as the source of the Group VIB metal, and one or more of nickel or cobalt salts as the source of the Group VIII metal. Further, in the impregnation solution containing Group VIB and Group VIII metals, the mass content of the Group VIB metal (calculated as oxide) is 3.0–30.0 g / 100 mL, and the mass content of the Group VIII metal (calculated as oxide) is 1.0–10.0 g / 100 mL.

[0035] Further, in step (3), the pH value of the impregnation solution containing Group VIB and Group VIII metals is 8.0 to 10.0. Further, there are no particular limitations on the pH adjuster used; conventional pH adjusters in the art can be used.

[0036] Furthermore, in step (4), the mass ratio of macroporous alumina loaded with Group VIB and Group VIII metals to modified ZSM-5 molecular sieve is 0.2 to 1.0.

[0037] Further, in step (4), the molding can be carried out using conventional methods in the art, such as extrusion molding. Conventional molding aids can be added during the molding process, such as at least one of extrusion aids, pectinic acids, and binders; the extrusion aid can be guar gum powder, the pectinic acid can be at least one of citric acid and nitric acid, preferably citric acid and nitric acid; the binder can be aluminum sol or silica sol prepared from porous alumina, preferably aluminum sol. The amount of molding aid added accounts for 1.0% to 5.0% of the mass of the hydroconversion catalyst.

[0038] Further, in step (4), after molding, the catalyst is dried and calcined to obtain the catalyst. The drying temperature is 100℃~150℃, and the drying time is 2~4h; the calcination temperature is 400℃~600℃, and the calcination time is 3~5h.

[0039] Compared with the prior art, the present invention has the following beneficial effects:

[0040] In the delayed coking light naphtha processing method of the present invention, the saturation and conversion of delayed coking light naphtha are achieved by combining silicon capture, hydrorefining and hydroconversion processes, resulting in a hydrocarbon stream product that is free of olefins and has an increased content of n-alkanes. When used as ethylene feedstock, it has a higher triene yield.

[0041] In this invention's delayed coking light naphtha processing method, the olefin content in the refined naphtha stream is controlled through silica capture and hydrorefining. Trace amounts of olefins act as initiators, rapidly reacting with the acidic centers of the catalyst to generate carbocations. These carbocations continue to propagate after the reaction, allowing the reaction to proceed continuously and thus improving the conversion capacity of the light naphtha conversion stage. The modified ZSM-5 molecular sieve used in this invention's hydroconversion catalyst has a higher mesopore content, improving the accessibility of acidic centers and promoting the contact and reaction between isomeric hydrocarbons in the feedstock and acidic centers, thereby increasing the conversion selectivity of isomeric hydrocarbons. The hydrogenation metal in the hydroconversion catalyst is mainly supported on macroporous alumina, resulting in low hydrogenation activity within the molecular sieve channels, reducing carbocation saturation, and promoting the light naphtha conversion reaction. The n-olefins or small molecule olefins obtained from the light hydrocarbon conversion diffuse out of the molecular sieve channels and rapidly contact the hydrogenation centers on the alumina, undergoing a saturation reaction and preventing carbon deposition. Using this processing method, the proportion of n-olefins in the product obtained from the light naphtha conversion can reach 59%. Detailed Implementation

[0042] The following examples and comparative examples further illustrate the role and effect of the technical solution of the present invention, but the following examples do not constitute a limitation on the scope of protection of the present invention.

[0043] In this invention, the total amount C of Group VIB and Group VIII metals on the catalyst surface is... ms The elemental composition and state of the catalyst surface (3 nm–10 nm) were determined by X-ray photoelectron spectroscopy (XPS) using a Thermofisher Multilab 2000 electron spectrometer. The excitation source was Mg Kα, and the cathode voltage and current were 13 kV and 20 mA, respectively. The electron binding energy was calibrated using C1s (284.6 eV). The data obtained directly from XPS are mass percentages, C ms The values ​​are converted to mol / g.

[0044] In this invention, the total content of Group VIB and Group VIII metals in the bulk phase of the catalyst is C. mtThe results were obtained by X-ray fluorescence spectroscopy (XRF) analysis using a ZSX100e X-ray fluorescence spectrometer. The spectral line was Kα, the crystal was LiF1, the target material was Rh, the detector was SC scintillation, the timing was 20 s, and the optical atmosphere was vacuum. The data obtained directly from XRF were mass percentages, C. mt The values ​​are converted to mol / g. The total amount of Group VIB and Group VIII metals in the bulk catalyst refers to the total amount of Group VIB and Group VIII metals in the entire catalyst.

[0045] In this invention, the specific surface area, pore volume, and pore distribution were measured using the following method: an ASAP 2420 cryogenic liquid nitrogen physical adsorption instrument manufactured by Microlithics, Inc. was used, with a pretreatment temperature of 300°C and a pretreatment time of 4 hours.

[0046] In this invention, the amount of carbon deposited on the catalyst is measured by an infrared sulfur and carbon analyzer. A high-frequency infrared carbon and sulfur analyzer of EMIA-820V from HORIBA Corporation of Japan is used. The high-frequency induction heating furnace has an output power of 2300W and a frequency of 20MHz. The sample heating temperature is 1800℃. The carbon content of the catalyst is calculated based on the peak area of ​​the vibrational spectrum of CO2 generated by the combustion of carbon deposits.

[0047] In this invention, unless otherwise specified, all percentages (%) in the embodiments and comparative examples refer to mass fractions.

[0048] In this invention, the ZSM-5 molecular sieve involved in the embodiments and comparative examples is a commercially available product with a pore volume of 0.18 cm³. 3 / g, specific surface area is 191cm³ 2 / g, the silicon-to-aluminum molar ratio of SiO2:Al2O3 is 55. The macroporous alumina is a commercially available product with a pore volume of 1.1 cm³. 3 / g, specific surface area is 350cm² 2 / g.

[0049] The silicon-scavenging agent involved in the embodiments and comparative examples of this invention is FHRS-3, whose properties are as follows: specific surface area 382 cm². 2 / g, pore volume 0.58cm 3 / g.

[0050] The hydrorefining catalyst involved in the embodiments and comparative examples of this invention is FF-66, a commercial catalyst developed by Sinopec (Dalian) Petrochemical Research Institute Co., Ltd.

[0051] In this invention, the conversion rate of light naphtha isohydrocarbons, the selectivity of isohydrocarbon conversion, and the proportion of n-alkanes in the products are calculated as follows:

[0052] Isomer hydrocarbon conversion rate = (mass of isomer C5 and C6 hydrocarbons in feedstock - mass of isomer C5 and C6 hydrocarbons in product) / (mass of isomer C5 and C6 hydrocarbons in feedstock) × 100%;

[0053] Isomer hydrocarbon conversion selectivity = (mass of isomer C5 and C6 hydrocarbons in feedstock - mass of isomer C5 and C6 hydrocarbons in product) / (mass of C5 and C6 hydrocarbons in feedstock - mass of C5 and C6 hydrocarbons in product) × 100%;

[0054] The proportion of n-alkane = mass of n-alkane in the product / total mass of the product × 100%, where n-alkane includes ethane, propane, n-butane, n-pentane and n-hexane.

[0055] In this invention, in each example, the nickel source used is nickel nitrate, and the tungsten source is ammonium metatungstate.

[0056] Example 1

[0057] Preparation of modified ZSM-5 molecular sieve (denoted as Z-1): Commercially available ZSM-5 molecular sieve was placed in a hydrothermal treatment furnace and treated at 450℃ and 0.1MPa steam pressure for 2 hours to obtain modified ZSM-5 molecular sieve Z-1, with a mesopore volume of 0.09 cm³. 3 / g, total pore volume is 0.17cm³ 3 / g, mesoporous pore volume accounts for 52% of the total pore volume.

[0058] 100g of macroporous alumina was impregnated for 2h with 200mL of an aqueous solution containing 2.0g / 100mL of nickel (calculated as oxide, nickel source was nickel nitrate), 10.0g / 100mL of tungsten (calculated as oxide, tungsten source was ammonium metatungstate), and pH 9.0. The impregnation pressure was 1.0MPa and the impregnation temperature was 60℃. After impregnation, the alumina was cooled to room temperature, filtered, dried at 120℃ for 3h, and calcined at 450℃ for 3h to obtain macroporous alumina loaded with nickel oxide and tungsten oxide, denoted as A-NW-1.

[0059] Take 100g of Z-1 molecular sieve and 30g of A-NW-1 material, mix them, add 85g of pectinic acid (concentration 1wt%) and 1g of extrusion aid guar gum powder, then crush, extrude into strips, dry at 120℃ for 3h, and calcine at 500℃ for 4h to obtain the catalyst, denoted as CAT-1, whose properties are shown in Table 1.

[0060] Example 2

[0061] Preparation of modified ZSM-5 molecular sieve (denoted as Z-2): Commercially available ZSM-5 molecular sieve was placed in a hydrothermal treatment furnace and treated at 500℃ and 0.1MPa steam pressure for 2 hours to obtain modified ZSM-5 molecular sieve Z-2, with a mesopore volume of 0.11 cm³. 3 / g, total pore volume is 0.17cm³ 3 / g, mesoporous pore volume accounts for 65% of the total pore volume.

[0062] The macroporous alumina loaded with nickel oxide and tungsten oxide in this example is the same as the macroporous alumina loaded with nickel oxide and tungsten oxide in Example 1 (denoted as A-NW-1);

[0063] Take 100g of Z-2 molecular sieve and 30g of A-NW-1 material, mix them, add 85g of pectinic acid (concentration 1wt%) and 1g of extrusion aid guar gum powder, then crush, extrude into strips, dry at 120℃ for 3h, and calcine at 500℃ for 4h to obtain the catalyst, denoted as CAT-2, whose properties are shown in Table 1.

[0064] Example 3

[0065] The modified ZSM-5 molecular sieve in this embodiment is the same as the modified ZSM-5 molecular sieve in Example 1 (denoted as Z-1).

[0066] Take 100g of macroporous alumina and immerse it in 200mL of an aqueous solution with a nickel content of 1.0g / 100mL, a tungsten content of 5.0g / 100mL, and a pH of 9.0 for 2h. The immersion pressure is 0.5MPa and the immersion temperature is 70℃. After immersion, cool to room temperature, filter, dry at 120℃ for 3h, and calcine at 450℃ for 3h to obtain macroporous alumina loaded with nickel oxide and tungsten oxide, denoted as A-NW-2.

[0067] Take 100g of Z-1 molecular sieve and 30g of A-NW-2 material, mix them, add 85g of pectinic acid (concentration 1wt%) and 1g of extrusion aid guar gum powder, then crush, extrude into strips, dry at 120℃ for 3h, and calcine at 500℃ for 4h to obtain the catalyst, denoted as CAT-3, whose properties are shown in Table 1.

[0068] Example 4

[0069] The modified ZSM-5 molecular sieve in this embodiment is the same as the modified ZSM-5 molecular sieve in Example 2 (denoted as Z-2).

[0070] Take 100g of macroporous alumina and immerse it in 200mL of an aqueous solution with a nickel content of 2.0g / 100mL, a tungsten content of 10.0g / 100mL, and a pH of 10.0 for 2h. The immersion pressure is 2.0MPa and the immersion temperature is 80℃. After immersion, cool to room temperature, filter, dry at 120℃ for 3h, and calcine at 450℃ for 3h to obtain macroporous alumina loaded with nickel oxide and tungsten oxide, denoted as A-NW-3.

[0071] Take 80g of Z-2 molecular sieve and 50g of A-NW-3 material, mix them, add 90g of pectinic acid (concentration 1wt%) and 1g of extrusion aid guar gum powder, then crush, extrude into strips, dry at 120℃ for 3h, and calcine at 500℃ for 4h to obtain the catalyst, denoted as CAT-4, whose properties are shown in Table 1.

[0072] Example 5

[0073] Preparation of modified ZSM-5 molecular sieve (denoted as Z-3): Commercially available ZSM-5 molecular sieve was placed in a hydrothermal treatment furnace and treated at 600℃ and 0.1MPa steam pressure for 2 hours to obtain modified ZSM-5 molecular sieve Z-3, with a mesopore volume of 0.14 cm³. 3 / g, total pore volume is 0.18cm³ 3 / g, mesoporous pore volume accounts for 78% of the total pore volume.

[0074] The macroporous alumina loaded with nickel oxide and tungsten oxide in this example is the same as the macroporous alumina loaded with nickel oxide and tungsten oxide in Example 1 (denoted as A-NW-1);

[0075] Take 100g of Z-3 molecular sieve and 30g of A-NW-1 material, mix them, add 85g of pectinic acid (concentration 1wt%) and 1g of extrusion aid guar gum powder, then crush, extrude into strips, dry at 120℃ for 3h, and calcine at 500℃ for 4h to obtain the catalyst, denoted as CAT-5, whose properties are shown in Table 1.

[0076] Example 6

[0077] The modified ZSM-5 molecular sieve in this embodiment is the same as the modified ZSM-5 molecular sieve in Example 2 (denoted as Z-2).

[0078] Take 100g of macroporous alumina and immerse it in 200mL of an aqueous solution with a nickel content of 1.0g / 100mL, a tungsten content of 5.0g / 100mL (based on oxides), and a pH of 8.0 for 2h. The immersion pressure is 2.0MPa and the immersion temperature is 80℃. After immersion, cool to room temperature, filter, dry at 120℃ for 3h, and calcine at 450℃ for 3h to obtain macroporous alumina loaded with nickel oxide and tungsten oxide, denoted as A-NW-4.

[0079] Take 50g of Z-2 molecular sieve and 50g of A-NW-4 material, mix them, add 75g of pectinic acid (concentration 1wt%) and 1g of extrusion aid guar gum powder, then crush, extrude into strips, dry at 120℃ for 3h, and calcine at 500℃ for 4h to obtain the catalyst, denoted as CAT-6, whose properties are shown in Table 1.

[0080] Comparative Example 1

[0081] 100g of modified ZSM-5 molecular sieve (denoted as Z-1) and 45g of macroporous alumina were mixed, and 85g of pectinic acid (1wt%) and 1g of extrusion aid guar gum powder were added. The mixture was then rolled, extruded into strips, dried at 120℃ for 3 hours, and calcined at 500℃ for 4 hours to obtain DA-NW-1. 100g of DA-NW-1 was impregnated in 200mL of an aqueous solution containing 1.5g / 100mL of nickel and 8.0g / 100mL of tungsten (based on oxides) for 2 hours. After filtration, drying at 120℃ for 3 hours, and calcining at 500℃ for 4 hours, a catalyst was obtained, denoted as DCAT-1. Its properties are shown in Table 1.

[0082] Comparative Example 2

[0083] The macroporous alumina loaded with nickel oxide and tungsten oxide in this example is the same as the macroporous alumina loaded with nickel oxide and tungsten oxide in Example 1 (denoted as A-NW-1).

[0084] Take 100g of commercially available ZSM-5 molecular sieve and 30g of A-NW-1 material, mix them together, add 85g of pectinic acid (concentration 1wt%) and 1g of extrusion aid guar gum powder, then crush, extrude into strips, dry at 120℃ for 3h, and calcine at 500℃ for 4h to obtain the catalyst, denoted as DCAT-2, whose properties are shown in Table 1.

[0085] Comparative Example 3

[0086] The modified ZSM-5 molecular sieve in this example is the same as the modified ZSM-5 molecular sieve in Example 1 (denoted as Z-1).

[0087] Take 100g of macroporous alumina and immerse it in 200mL of an aqueous solution with a nickel content of 8.0g / 100mL, a tungsten content of 40.0g / 100mL (based on oxides), and a pH of 9.0 for 2h. The immersion pressure is 1.0MPa and the immersion temperature is 60℃. After immersion, cool to room temperature, filter, dry at 120℃ for 3h, and calcine at 450℃ for 3h to obtain macroporous alumina loaded with nickel oxide and tungsten oxide, denoted as A-NW-D3.

[0088] Take 100g of Z-1 molecular sieve and 30g of A-NW-D3 material, mix them, add 85g of pectinic acid (concentration 1wt%) and 1g of extrusion aid guar gum powder, then crush, extrude into strips, dry at 120℃ for 3h, and calcine at 500℃ for 4h to obtain the catalyst, denoted as DCAT-3, whose properties are shown in Table 1.

[0089] Comparative Example 4

[0090] The modified ZSM-5 molecular sieve in this example is the same as the modified ZSM-5 molecular sieve in Example 1 (denoted as Z-1).

[0091] Take 100g of macroporous alumina and immerse it in 200mL of an aqueous solution (pH 5.4) containing 2.0g / 100mL of nickel and 10.0g / 100mL of tungsten (calculated as oxides) for 2h. After filtration, drying at 120℃ for 3h and calcining at 450℃ for 3h, macroporous alumina loaded with nickel oxide and tungsten oxide is obtained, denoted as A-NW-D4.

[0092] Take 100g of Z-1 molecular sieve and 30g of A-NW-D4 material, mix them, add 85g of pectinic acid (concentration 1wt%) and 1g of extrusion aid guar gum powder, then crush, extrude into strips, dry at 120℃ for 3h, and calcine at 500℃ for 4h to obtain the catalyst, denoted as DCAT-4, whose properties are shown in Table 1.

[0093] Comparative Example 5

[0094] The modified ZSM-5 molecular sieve in this example is the same as the modified ZSM-5 molecular sieve in Example 1 (denoted as Z-1).

[0095] Take 100g of small-pore alumina (pore volume 0.48cm³). 3 / g, specific surface area is 251cm³ 2 The nickel oxide and tungsten oxide were impregnated in 200 mL of an aqueous solution with a nickel content of 2.0 g / 100 mL, a tungsten content of 10.0 g / 100 mL (based on oxides), and a pH of 9.0 for 2 h. The impregnation pressure was 1.0 MPa and the impregnation temperature was 60 °C. After impregnation, the mixture was cooled to room temperature, filtered, dried at 120 °C for 3 h, and calcined at 450 °C for 3 h to obtain a porous alumina loaded with nickel oxide and tungsten oxide, denoted as A-NW-D5.

[0096] Take 100g of Z-1 molecular sieve and 30g of A-NW-D5 material, mix them, add 85g of pectinic acid (concentration 1wt%) and 1g of extrusion aid guar gum powder, then crush, extrude into strips, dry at 120℃ for 3h, and calcine at 500℃ for 4h to obtain the catalyst, denoted as DCAT-5, whose properties are shown in Table 1.

[0097] Table 1. Composition and properties of catalysts in the examples and comparative examples.

[0098]

[0099] Note: The alumina content of DCAT-5 is calculated based on microporous alumina, while the content of the other catalysts is calculated based on macroporous alumina.

[0100] Example 7

[0101] 0.2 mL of silica scavenger FHRS-3 and 0.8 mL of catalyst FF-66 were placed in the upper and lower parts of the first reactor of the fixed-bed reactor, respectively. 3.0 mL of catalyst CAT-1 was placed in the second reactor of the fixed-bed reactor. Delayed coking light naphtha was introduced into the first reactor at a flow rate of 10 mL / h to contact the silica scavenger FHRS-3 and catalyst FF-66 for the first reaction. The resulting first-stage effluent was introduced into the second reactor to contact the catalyst CAT-1 for the second reaction. The reaction conditions for the first reaction were as follows: reaction pressure 4.0 MPa, hydrogen-to-oil volume ratio 400:1, and reaction temperature 150℃. The reaction conditions for the second reaction were as follows: reaction pressure 4.0 MPa, hydrogen-to-oil volume ratio 400:1, and reaction temperature 400℃. The properties of the feedstock are shown in Table 2. After 160 h of operation, the bromine value, isomeric hydrocarbon conversion rate, isomeric hydrocarbon conversion selectivity, n-hydrocarbon content of the product, and catalyst carbon deposition of the first-stage effluent are shown in Table 3.

[0102] Example 8

[0103] 0.2 mL of silica scavenger FHRS-3 and 0.8 mL of catalyst FF-66 were placed in the upper and lower parts of the first reactor of the fixed-bed reactor, respectively. 8.0 mL of catalyst CAT-2 was placed in the second reactor of the fixed-bed reactor. Delayed coking light naphtha was introduced into the first reactor at a flow rate of 10 mL / h to contact the silica scavenger FHRS-3 and catalyst FF-66 for the first reaction. The resulting first-stage effluent was then introduced into the second reactor to contact the catalyst CAT-2 for the second reaction. The reaction conditions for the first reaction were as follows: reaction pressure 4.0 MPa, hydrogen-to-oil volume ratio 400:1, and reaction temperature 150℃. The reaction conditions for the second reaction were as follows: reaction pressure 4.0 MPa, hydrogen-to-oil volume ratio 400:1, and reaction temperature 400℃. The properties of the feedstock are shown in Table 2. After 160 h of operation, the bromine value, isomeric hydrocarbon conversion rate, isomeric hydrocarbon conversion selectivity, n-hydrocarbon content of the product, and catalyst carbon deposition of the first-stage effluent are shown in Table 3.

[0104] Example 9

[0105] 0.2 mL of silica scavenger FHRS-3 and 0.8 mL of catalyst FF-66 were placed in the upper and lower parts of the first reactor of the fixed-bed reactor, respectively. 5.0 mL of catalyst CAT-3 was placed in the second reactor of the fixed-bed reactor. Delayed coking light naphtha was introduced into the first reactor at a flow rate of 10 mL / h to contact the silica scavenger FHRS-3 and catalyst FF-66 for the first reaction. The resulting first-stage effluent was introduced into the second reactor to contact the catalyst CAT-3 for the second reaction. The reaction conditions for the first reaction were as follows: reaction pressure 4.0 MPa, hydrogen-to-oil volume ratio 400:1, and reaction temperature 150℃. The reaction conditions for the second reaction were as follows: reaction pressure 4.0 MPa, hydrogen-to-oil volume ratio 400:1, and reaction temperature 400℃. The properties of the feedstock are shown in Table 2. After 160 h of operation, the bromine value, isomeric hydrocarbon conversion rate, isomeric hydrocarbon conversion selectivity, n-hydrocarbon content of the product, and catalyst carbon deposition of the first-stage effluent are shown in Table 3.

[0106] Example 10

[0107] 0.2 mL of silica scavenger FHRS-3 and 0.8 mL of catalyst FF-66 were placed in the upper and lower parts of the first reactor of the fixed-bed reactor, respectively. 5.0 mL of catalyst CAT-4 was placed in the second reactor of the fixed-bed reactor. Delayed coking light naphtha was introduced into the first reactor at a flow rate of 10 mL / h to contact the silica scavenger FHRS-3 and catalyst FF-66 for the first reaction. The resulting first-stage effluent was introduced into the second reactor to contact the catalyst CAT-4 for the second reaction. The reaction conditions for the first reaction were as follows: reaction pressure 4.0 MPa, hydrogen-to-oil volume ratio 400:1, and reaction temperature 150℃. The reaction conditions for the second reaction were as follows: reaction pressure 4.0 MPa, hydrogen-to-oil volume ratio 400:1, and reaction temperature 400℃. The properties of the feedstock are shown in Table 2. After 160 h of operation, the bromine value, isomeric hydrocarbon conversion rate, isomeric hydrocarbon conversion selectivity, n-hydrocarbon content of the product, and catalyst carbon deposition of the first-stage effluent are shown in Table 3.

[0108] Example 11

[0109] 0.2 mL of silica scavenger FHRS-3 and 0.8 mL of catalyst FF-66 were placed in the upper and lower parts of the first reactor of the fixed-bed reactor, respectively. 5.0 mL of catalyst CAT-5 was placed in the second reactor of the fixed-bed reactor. Delayed coking light naphtha was introduced into the first reactor at a flow rate of 10 mL / h to contact the silica scavenger FHRS-3 and catalyst FF-66 for the first reaction. The resulting first-stage effluent was introduced into the second reactor to contact the catalyst CAT-5 for the second reaction. The reaction conditions for the first reaction were as follows: reaction pressure 4.0 MPa, hydrogen-to-oil volume ratio 400:1, and reaction temperature 150℃. The reaction conditions for the second reaction were as follows: reaction pressure 4.0 MPa, hydrogen-to-oil volume ratio 400:1, and reaction temperature 400℃. The properties of the feedstock are shown in Table 2. After 160 h of operation, the bromine value, isomeric hydrocarbon conversion rate, isomeric hydrocarbon conversion selectivity, n-hydrocarbon content of the product, and catalyst carbon deposition of the first-stage effluent are shown in Table 3.

[0110] Example 12

[0111] 0.1 mL of silica scavenger FHRS-3 and 1.0 mL of catalyst FF-66 were placed in the upper and lower parts of the first reactor of the fixed-bed reactor, respectively. 5.0 mL of catalyst CAT-6 was placed in the second reactor of the fixed-bed reactor. Delayed coking light naphtha was introduced into the first reactor at a flow rate of 10 mL / h to contact the silica scavenger FHRS-3 and catalyst FF-66 for the first reaction. The resulting first-stage effluent was introduced into the second reactor to contact the catalyst CAT-6 for the second reaction. The reaction conditions for the first reaction were as follows: reaction pressure 4.0 MPa, hydrogen-to-oil volume ratio 400:1, and reaction temperature 150℃. The reaction conditions for the second reaction were as follows: reaction pressure 4.0 MPa, hydrogen-to-oil volume ratio 400:1, and reaction temperature 400℃. The properties of the feedstock are shown in Table 2. After 160 h of operation, the bromine value, isomeric hydrocarbon conversion rate, isomeric hydrocarbon conversion selectivity, n-hydrocarbon content of the product, and catalyst carbon deposition of the first-stage effluent are shown in Table 3.

[0112] Comparative Example 6

[0113] 1.0 mL of catalyst FF-66 was placed in the first reactor of the fixed-bed reactor, and 5.0 mL of catalyst CAT-3 was placed in the second reactor. Delayed coking light naphtha was introduced into the first reactor at a flow rate of 10 mL / h to contact with catalyst FF-66 for the first reaction. The resulting first-stage effluent was then introduced into the second reactor to contact with catalyst CAT-3 for the second reaction. The reaction conditions for the first reaction were as follows: reaction pressure 4.0 MPa, hydrogen-to-oil volume ratio 400:1, and reaction temperature 150℃. The reaction conditions for the second reaction were as follows: reaction pressure 4.0 MPa, hydrogen-to-oil volume ratio 400:1, and reaction temperature 400℃. The properties of the feedstock are shown in Table 2. After 160 h of operation, the bromine value, isomeric hydrocarbon conversion rate, isomeric hydrocarbon conversion selectivity, n-hydrocarbon content of the product, and catalyst carbon deposition of the first-stage effluent are shown in Table 4.

[0114] Comparative Example 7

[0115] In the first reactor, no catalyst was loaded. 5.0 mL of catalyst CAT-3 was placed in the second reactor of the fixed-bed reactor. Delayed coking light naphtha was introduced into the second reactor at a flow rate of 10 mL / h to contact the catalyst CAT-3. The reaction conditions in the second reactor were as follows: reaction pressure 4.0 MPa, hydrogen-to-oil volume ratio 150:1, and liquid hourly space velocity (LHSV) 1.0 h⁻¹. -1 The reaction temperature was 400℃. The properties of the raw materials are shown in Table 2. After running for 160 hours, the bromine value, isomeric hydrocarbon conversion rate, isomeric hydrocarbon conversion selectivity, n-hydrocarbon content of the product, and catalyst carbon deposition of the first reaction effluent are shown in Table 4.

[0116] Comparative Example 8

[0117] 0.5 mL of silica scavenger FHRS-3 and 1.5 mL of catalyst FF-66 were placed in the upper and lower parts of the first reactor of the fixed-bed reactor, respectively. 5.0 mL of catalyst CAT-3 was placed in the second reactor of the fixed-bed reactor. Delayed coking light naphtha was introduced into the first reactor at a flow rate of 10 mL / h to contact the silica scavenger FHRS-3 and catalyst FF-66 for the first reaction. The resulting first-stage effluent was introduced into the second reactor to contact the catalyst CAT-3 for the second reaction. The reaction conditions for the first reaction were as follows: reaction pressure 4.0 MPa, hydrogen-to-oil volume ratio 400:1, and reaction temperature 150℃. The reaction conditions for the second reaction were as follows: reaction pressure 4.0 MPa, hydrogen-to-oil volume ratio 400:1, and reaction temperature 400℃. The properties of the feedstock are shown in Table 2. After 160 h of operation, the bromine value, isomeric hydrocarbon conversion rate, isomeric hydrocarbon conversion selectivity, n-hydrocarbon content of the product, and catalyst carbon deposition of the first-stage effluent are shown in Table 4.

[0118] Comparative Example 9

[0119] 0.2 mL of silica scavenger FHRS-3 and 0.8 mL of catalyst FF-66 were placed in the upper and lower parts of the first reactor of the fixed-bed reactor, respectively. 5.0 mL of catalyst DCAT-1 was placed in the second reactor of the fixed-bed reactor. Delayed coking light naphtha was introduced into the first reactor at a flow rate of 10 mL / h to contact the silica scavenger FHRS-3 and catalyst FF-66 for the first reaction. The resulting first-stage effluent was introduced into the second reactor to contact the catalyst DCAT-1 for the second reaction. The reaction conditions for the first reaction were as follows: reaction pressure 4.0 MPa, hydrogen-to-oil volume ratio 400:1, and reaction temperature 150℃. The reaction conditions for the second reaction were as follows: reaction pressure 4.0 MPa, hydrogen-to-oil volume ratio 400:1, and reaction temperature 400℃. The properties of the feedstock are shown in Table 2. After 160 h of operation, the bromine value, isomeric hydrocarbon conversion rate, isomeric hydrocarbon conversion selectivity, n-hydrocarbon content of the product, and catalyst carbon deposition of the first-stage effluent are shown in Table 4.

[0120] Comparative Example 10

[0121] 0.2 mL of silica scavenger FHRS-3 and 0.8 mL of catalyst FF-66 were placed in the upper and lower parts of the first reactor of the fixed-bed reactor, respectively. 5.0 mL of catalyst DCAT-2 was placed in the second reactor of the fixed-bed reactor. Delayed coking light naphtha was introduced into the first reactor at a flow rate of 10 mL / h to contact the silica scavenger FHRS-3 and catalyst FF-66 for the first reaction. The resulting first-stage effluent was introduced into the second reactor to contact the catalyst DCAT-2 for the second reaction. The reaction conditions for the first reaction were as follows: reaction pressure 4.0 MPa, hydrogen-to-oil volume ratio 400:1, and reaction temperature 150℃. The reaction conditions for the second reaction were as follows: reaction pressure 4.0 MPa, hydrogen-to-oil volume ratio 400:1, and reaction temperature 400℃. The properties of the feedstock are shown in Table 2. After 160 h of operation, the bromine value, isomeric hydrocarbon conversion rate, isomeric hydrocarbon conversion selectivity, n-hydrocarbon content of the product, and catalyst carbon deposition of the first-stage effluent are shown in Table 4.

[0122] Comparative Example 11

[0123] 0.2 mL of silica scavenger FHRS-3 and 0.8 mL of catalyst FF-66 were placed in the upper and lower parts of the first reactor of the fixed-bed reactor, respectively. 5.0 mL of catalyst DCAT-3 was placed in the second reactor of the fixed-bed reactor. Delayed coking light naphtha was introduced into the first reactor at a flow rate of 10 mL / h to contact the silica scavenger FHRS-3 and catalyst FF-66 for the first reaction. The resulting first-stage effluent was introduced into the second reactor to contact the catalyst DCAT-3 for the second reaction. The reaction conditions for the first reaction were as follows: reaction pressure 4.0 MPa, hydrogen-to-oil volume ratio 400:1, and reaction temperature 150℃. The reaction conditions for the second reaction were as follows: reaction pressure 4.0 MPa, hydrogen-to-oil volume ratio 400:1, and reaction temperature 400℃. The properties of the feedstock are shown in Table 2. After 160 h of operation, the bromine value, isomeric hydrocarbon conversion rate, isomeric hydrocarbon conversion selectivity, n-hydrocarbon content of the product, and catalyst carbon deposition of the first-stage effluent are shown in Table 4.

[0124] Comparative Example 12

[0125] 0.2 mL of silica scavenger FHRS-3 and 0.8 mL of catalyst FF-66 were placed in the upper and lower parts of the first reactor of the fixed-bed reactor, respectively. 5.0 mL of catalyst DCAT-4 was placed in the second reactor of the fixed-bed reactor. Delayed coking light naphtha was introduced into the first reactor at a flow rate of 10 mL / h to contact the silica scavenger FHRS-3 and catalyst FF-66 for the first reaction. The resulting first-stage effluent was introduced into the second reactor to contact the catalyst DCAT-4 for the second reaction. The reaction conditions for the first reaction were as follows: reaction pressure 4.0 MPa, hydrogen-to-oil volume ratio 400:1, and reaction temperature 150℃. The reaction conditions for the second reaction were as follows: reaction pressure 4.0 MPa, hydrogen-to-oil volume ratio 400:1, and reaction temperature 400℃. The properties of the feedstock are shown in Table 2. After 160 h of operation, the bromine value, isomeric hydrocarbon conversion rate, isomeric hydrocarbon conversion selectivity, n-hydrocarbon content of the product, and catalyst carbon deposition of the first-stage effluent are shown in Table 4.

[0126] Comparative Example 13

[0127] 0.2 mL of silica scavenger FHRS-3 and 0.8 mL of catalyst FF-66 were placed in the upper and lower parts of the first reactor of the fixed-bed reactor, respectively. 5.0 mL of catalyst DCAT-5 was placed in the second reactor of the fixed-bed reactor. Delayed coking light naphtha was introduced into the first reactor at a flow rate of 10 mL / h to contact the silica scavenger FHRS-3 and catalyst FF-66 for the first reaction. The resulting first-stage effluent was introduced into the second reactor to contact the catalyst DCAT-5 for the second reaction. The reaction conditions for the first reaction were as follows: reaction pressure 4.0 MPa, hydrogen-to-oil volume ratio 400:1, and reaction temperature 150℃. The reaction conditions for the second reaction were as follows: reaction pressure 4.0 MPa, hydrogen-to-oil volume ratio 400:1, and reaction temperature 400℃. The properties of the feedstock are shown in Table 2. After 160 h of operation, the bromine value, isomeric hydrocarbon conversion rate, isomeric hydrocarbon conversion selectivity, n-hydrocarbon content of the product, and catalyst carbon deposition of the first-stage effluent are shown in Table 4.

[0128] Table 2 Composition and properties of delayed coking light naphtha

[0129] composition content n-hydrocarbons 44.5wt% Isomers 41.6wt% <![CDATA[C5-C6 isoparaffin]]> 40.3wt% Cycloalkanes 5.9wt% Aromatics 0.3wt% bromine value 66.4gBr / 100g Initial boiling point 28.9℃ Final boiling point 71.3℃

[0130] Table 3 Evaluation results of the catalysts in the examples

[0131]

[0132] Table 4 Evaluation results of the comparative example catalysts

[0133]

[0134] The specific embodiments of the present invention have been described in detail above; however, the present invention is not limited thereto. Within the scope of the inventive concept, various simple modifications can be made to the technical solutions of the present invention, including combining the various technical features in any other suitable manner. These simple modifications and combinations should also be considered as the content disclosed in the present invention and are all within the protection scope of the present invention.

Claims

1. A method for processing delayed coking light naphtha, comprising: (I) Delayed coking light naphtha is mixed with hydrogen and then enters the hydrorefining reaction zone for hydrorefining reaction to obtain refined naphtha stream. The hydrorefining zone is sequentially filled with silica scavenger and hydrorefining catalyst along the material flow direction. (II) The refined naphtha stream from step (I) enters the hydroconversion reaction zone for hydroconversion reaction to obtain the converted delayed coking light naphtha.

2. The method according to claim 1, characterized in that: The properties of the delayed coking light naphtha are as follows: initial boiling point is 15-30℃, and final boiling point is 60-90℃; And / or, in the delayed coking light naphtha, the silicon content (SiO2) is 5-20 μg / g and the bromine value is 30-70 gBr / 100g; And / or, the C5-C6 isomeric hydrocarbon content in the delayed coking light naphtha is 20wt% to 50wt%, preferably 30wt% to 50wt%, based on a total weight of 100wt% of the delayed coking light naphtha.

3. The method according to claim 1, characterized in that: In step (I), the bromine value of the refined naphtha stream is 0.08–0.30 gBr / 100 g.

4. The method according to claim 1, characterized in that: In step (I), the volume ratio of the silicon-collecting agent to the hydrorefining catalyst is 1:4 to 1:

10.

5. The method according to claim 1, characterized in that: In step (I), the hydrorefining conditions are: reaction temperature 100–250°C, reaction pressure 2–8 MPa, and liquid hourly space velocity (LHSV) 8.0–15.0 h⁻¹. -1 The hydrogen-to-oil volume ratio is 100:1 to 2000:1; And / or, in step (II), the hydrogenation conversion conditions are: reaction temperature 300–450°C, reaction pressure 2–8 MPa, and liquid hourly space velocity 0.1–5.0 h⁻¹. -1 The hydrogen-to-oil volume ratio is 100:1 to 2000:

1.

6. The method according to claim 1, characterized in that: In step (II), the hydroconversion reaction zone is filled with a hydroconversion catalyst; preferably, the volume ratio of the hydroconversion catalyst to the hydrorefining catalyst is 3:1 to 10:

1.

7. The method according to claim 7, characterized in that: The hydroconversion catalyst comprises modified ZSM-5 molecular sieve and macroporous alumina supported on Group VIB and Group VIII metals.

8. The method according to claim 6 or 7, characterized in that: Based on the weight of the hydroconversion catalyst, the modified ZSM-5 molecular sieve content is 35wt% to 80wt%, the macroporous alumina content is 15wt% to 50wt%, and the total content of Group VIB and Group VIII metals as oxides is 0.5wt% to 8.0wt%. Preferably, based on the weight of the hydroconversion catalyst, the content of modified ZSM-5 molecular sieve is 50wt% to 75wt%, the content of macroporous alumina is 20wt% to 30wt%, and the total content of Group VIB and Group VIII metals as oxides is 1.0wt% to 5.0wt%. Preferably, in the macroporous alumina loaded with Group VIB and Group VIII metals, the mass ratio of the Group VIB metal (calculated as oxide) to the Group VIII metal (calculated as oxide) is 2.0 to 6.

0.

9. The method according to claim 7, characterized in that: In the hydrogenation conversion catalyst, the mesoporous pore volume of the modified ZSM-5 molecular sieve accounts for 50% to 80% of the total pore volume, preferably 60% to 70%.

10. The method according to claim 7, characterized in that: The total content C of Group VIB and Group VIII metals on the surface of the hydroconversion catalyst ms The total content of Group VIB and Group VIII metals in the catalyst bulk phase, C mt The molar ratio is 0.60 to 0.90.

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