A hydrocracking method for heavy and inferior wax oil

Abstract

The present invention relates to the technical field of heavy and low-quality wax oil processing, and discloses a hydrocracking method for heavy and low-quality wax oil. The method comprises: introducing the heavy and low-quality wax oil into a hydrorefining reaction zone I filled with a hydrorefining catalyst I to perform a first reaction to obtain a reaction effluent I; introducing the reaction effluent I into a hydrorefining reaction zone II filled with a hydrorefining catalyst II to perform a second reaction to obtain a reaction effluent II; introducing the reaction effluent II into a hydrocracking reaction zone filled with a hydrocracking catalyst to perform a third reaction to obtain a reaction effluent III; and separating the reaction effluent III to obtain a light naphtha fraction, a heavy naphtha fraction, a middle distillate oil, and a tail oil fraction. The method provided by the present invention can convert heavy and low-quality wax oil feedstock into clean fuel and high-quality chemical feedstock.

Description

Technical Field

[0001] The present invention relates to the technical field of heavy and inferior wax oil processing, and in particular to a hydrocracking method for heavy and inferior wax oil. Background Art

[0002] In recent years, with the continued development of the international energy and chemical industries, crude oil production has increased annually. However, as a non-renewable resource, the supply of high-quality and conventional crude oil has decreased and its price has risen with increasing production. Meanwhile, the supply of heavy and low-quality crude oil has increased annually, and its price has remained relatively low. Although heavy and low-quality crude oil is difficult and costly to process, refineries can achieve significant economic returns by selecting appropriate refining technologies and product solutions, processing this inexpensive and abundant source of crude oil.

[0003] To meet the processing needs of heavy and inferior crude oil, delayed coking technology for lightening heavy oil, as well as slurry bed, ebullating bed, and suspension bed heavy oil hydrocracking technologies have been further developed in recent years. However, heavy oil delayed coking, slurry bed, ebullating bed, and suspension bed heavy oil hydrocracking technologies will produce a considerable proportion of wax oil fractions or heavy distillate oils during the process of lightening residual oil or heavy oil. Compared with the wax oil raw materials obtained by conventional crude oil fractionation, the heavy and inferior wax oil fractions or wide-fraction heavy and inferior wax oils obtained by lightening heavy and inferior crude oil have the characteristics of high density, high dry point, high sulfur content, high nitrogen content, and high carbon residue value. The density at 20°C is about 0.94g / cm 3 -0.99g / cm 3 The distillation dry point is about 600℃-720℃, the aromatic content is about 60%-85%, the sulfur content is about 1.5%-3.5%, the nitrogen content is about 1800μg / g-5000μg / g, the residual carbon value is about 1.5%-5.0%, etc. It is difficult to realize the conversion of heavy and inferior wax oil raw materials into light products and high-value products under the existing hydrocracking technology.

[0004] CN104611051A and CN103102957A disclose a two-stage hydrocracking method for high-drying-point heavy distillate oil. This method uses a two-stage denitrogenation method to process heavy straight-run wax oil raw materials with low nitrogen content but complex nitrogen-containing compound structures to produce conventional hydrocracking products.

[0005] CN104611053A discloses a high-drying-point raw material hydrocracking process, which adopts a two-stage denitrification method and arranges a gas-liquid countercurrent contact reaction zone and a gas-liquid separation zone in one stage to realize the processing of heavy and inferior deep-drawn wax oil raw materials with a final distillation point of about 580°C-620°C to produce chemical raw materials and conventional clean fuel oil.

[0006] CN1940030A discloses a two-stage hydrocracking method for producing more diesel from high-nitrogen heavy crude oil. The method realizes the conversion of high-nitrogen heavy low-quality raw materials into high-quality diesel products by partitioning the feed and setting a hot flash tank between the first refining stage and the second refining stage.

[0007] CN103059944A discloses a hydrocracking process for processing inferior raw materials. The method adopts a two-stage hydrocracking process equipped with two fractionation units to convert inferior raw oil with a nitrogen content greater than 2000 μg / g and a final distillation point of approximately 450°C-620°C into chemical raw materials and conventional clean fuel oil.

[0008] CN103443251A discloses a hydrocracking method with feed bottom logistics treatment, which realizes the conversion of heavy and inferior raw materials by adopting a combined process of adsorption denitrogenation and removal of polynuclear aromatics and hydrocracking.

[0009] CN104560169A discloses a hydrocracking method for producing heavy naphtha from a high-nitrogen feedstock. The method realizes the conversion of the high-nitrogen feedstock into a heavy naphtha product by setting a second tail oil fraction conversion reaction zone.

[0010] CN104611010A and CN104611016A disclose a two-stage hydrocracking method for high-drying-point raw materials. This method realizes the conversion of heavy and low-quality raw materials with a final distillation point of 580°C-620°C into conventional clean oil products and chemical raw materials by setting up a gas-liquid separation method in the second-stage refining reaction zone equipped with a bulk catalyst.

[0011] CN104611022A discloses a two-stage hydrocracking method for inferior heavy distillate oil. The method achieves the conversion of heavy and inferior raw materials with a nitrogen content of 2000μg / g-15000μg / g and a final distillation point of 470°C-550°C into conventional clean fuels and chemical fuels by loading refined catalysts with different average pore sizes in two-stage hydrorefining reaction zones.

[0012] CN1508232A discloses a full-cycle hydrocracking process, which achieves the conversion of heavy crude oil with a distillation range of 200°C-600°C by adopting a hydrocracking method of first cracking and then refining, and using two or more high-nitrogen-resistant hydrocracking catalysts in the cracking reaction zone.

[0013] However, there are two main problems in the above existing hydrocracking technology:

[0014] On the one hand, the raw materials processed by existing hydrocracking technology are not of low quality and cannot meet the processing needs of low-quality raw materials; on the other hand, existing hydrocracking products mainly focus on processing conventional raw materials to produce clean fuel oil (diesel and jet fuel) and some chemical raw materials, and there is little research on the conversion of heavy and low-quality wax oil and product route development.

[0015] To this end, the development of hydrocracking technology to lighten and enhance the value of heavy and inferior wax oil raw materials is of great significance for refining and chemical companies to achieve high-value conversion of heavy and inferior crude oil. Summary of the Invention

[0016] The purpose of the present invention is to solve the problems of difficulty in converting heavy and low-quality wax oil raw materials and relatively single product routes in the prior art, and to provide a method for converting heavy and low-quality wax oil into clean fuel and high-quality chemical raw materials.

[0017] To achieve the above-mentioned object, the present invention provides a method for hydrocracking heavy low-quality wax oil, which is carried out in a device comprising a hydrorefining reaction zone and a hydrocracking reaction zone, wherein the hydrorefining reaction zone comprises a hydrorefining reaction zone I and a hydrorefining reaction zone II. The method comprises:

[0018] (1) introducing the heavy low-quality wax oil into the hydrotreating reaction zone I filled with the hydrotreating catalyst I to perform a first reaction to obtain a reaction effluent I;

[0019] (2) introducing the reaction effluent I into the hydrotreating reaction zone II filled with a hydrotreating catalyst II to perform a second reaction to obtain a reaction effluent II;

[0020] (3) introducing the reaction effluent II into the hydrocracking reaction zone filled with a hydrocracking catalyst to perform a third reaction to obtain a reaction effluent III;

[0021] (4) separating the reaction effluent III to obtain a light naphtha fraction, a heavy naphtha fraction, a middle distillate oil, and a tail oil fraction;

[0022] The heavy low-quality wax oil has a 95% distillation temperature of 580° C. to 700° C. according to ASTM D-1160, a nitrogen content of 1800 μg / g to 3500 μg / g, and a sum of two-ring or higher cycloalkanes and two-ring or higher aromatics of 50 wt % to 80 wt %.

[0023] Controlling the conditions of the first reaction so that the saturation rate of three-ring or higher aromatic hydrocarbons in the reaction effluent I is 75%-85%;

[0024] The reaction temperature of the first reaction is 10° C. to 30° C. lower than the reaction temperature of the second reaction.

[0025] The present invention sequentially introduces heavy and low-quality wax oil into a hydrorefining reaction zone I and a hydrorefining reaction zone II with a reaction temperature difference of no more than 10°C-30°C, performs a two-stage hydrorefining reaction, and coordinates the process conditions of the hydrorefining reaction to convert the heavy and low-quality wax oil raw material into clean fuel and high-quality chemical raw material. DETAILED DESCRIPTION

[0026] The endpoints of the ranges and any values disclosed herein are not limited to the precise ranges or values, and these ranges or values should be understood to include values close to these ranges or values. For numerical ranges, the endpoints of each range, the endpoints of each range and individual point values, and the individual point values can be combined with each other to obtain one or more new numerical ranges, which should be considered to be specifically disclosed herein.

[0027] In the present invention, the content of two-ring or higher cycloalkanes in the heavy low-quality wax oil = the mass of two-ring or higher cycloalkanes in the heavy low-quality wax oil / the total mass of hydrocarbon groups in the heavy low-quality wax oil × 100%.

[0028] In the present invention, the content of two-ring or higher aromatic hydrocarbons in the heavy low-quality wax oil = the mass of two-ring or higher aromatic hydrocarbons in the heavy low-quality wax oil / the total mass of hydrocarbon groups in the heavy low-quality wax oil × 100%.

[0029] In the present invention, the hydrocarbon family includes alkanes, cycloalkanes and aromatic hydrocarbons, and the alkanes include straight-chain alkanes and branched-chain alkanes.

[0030] In the present invention, the saturation rate of two-ring or higher aromatic hydrocarbons in the reaction effluent II = (the content of two-ring or higher aromatic hydrocarbons in the heavy inferior wax oil - the content of two-ring or higher aromatic hydrocarbons in the reaction effluent II) / the content of two-ring or higher aromatic hydrocarbons in the heavy inferior wax oil × 100%.

[0031] In the present invention, the saturation rate of three-ring or higher aromatic hydrocarbons in the reaction effluent I = (content of three-ring or higher aromatic hydrocarbons in the heavy inferior wax oil - content of three-ring or higher aromatic hydrocarbons in the reaction effluent I) / content of three-ring or higher aromatic hydrocarbons in the heavy inferior wax oil × 100%.

[0032] In the present invention, the yield of the tail oil fraction=the weight of the tail oil fraction / the weight of the heavy inferior wax oil×100%.

[0033] In the present invention, the 95% distillation temperature refers to the temperature corresponding to the moment when 95 wt% of the substances in the heavy inferior wax oil evaporate.

[0034] In the present invention, the pore size test is carried out with reference to the catalyst pore size distribution calculation method (nitrogen desorption isotherm calculation method) in SH / T 0572-1993, and the pore volume is tested with reference to the zeolite specific surface area and micropore volume determination method in NB / SH / T 0571-2021 catalyst.

[0035] In the present invention, the mesopore area and specific surface area are tested with reference to NB / SH / T 0571-2021 Method for Determining Specific Surface Area and Micropore Volume of Zeolite in Catalysts and GB / T 5816-1995 Method for Determining Surface Area of Catalysts and Adsorbents.

[0036] As described above, the present invention provides a method for hydrocracking heavy low-quality wax oil, which is carried out in a device comprising a hydrorefining reaction zone and a hydrocracking reaction zone, wherein the hydrorefining reaction zone comprises a hydrorefining reaction zone I and a hydrorefining reaction zone II. The method comprises:

[0037] (1) introducing the heavy low-quality wax oil into the hydrotreating reaction zone I filled with a hydrotreating catalyst I to perform a first reaction to obtain a reaction effluent I;

[0038] (2) introducing the reaction effluent I into the hydrotreating reaction zone II filled with a hydrotreating catalyst II to perform a second reaction to obtain a reaction effluent II;

[0039] (3) introducing the reaction effluent II into the hydrocracking reaction zone filled with a hydrocracking catalyst to perform a third reaction to obtain a reaction effluent III;

[0040] (4) separating the reaction effluent III to obtain a light naphtha fraction, a heavy naphtha fraction, a middle distillate oil, and a tail oil fraction;

[0041] The heavy low-quality wax oil has a 95% distillation temperature of 580° C. to 700° C. according to ASTM D-1160, a nitrogen content of 1800 μg / g to 3500 μg / g, and a sum of two-ring or higher cycloalkanes and two-ring or higher aromatics of 50 wt % to 80 wt %.

[0042] Controlling the conditions of the first reaction so that the saturation rate of three-ring or higher aromatic hydrocarbons in the reaction effluent I is 65%-85%;

[0043] The reaction temperature of the first reaction is 10° C. to 30° C. lower than the reaction temperature of the second reaction.

[0044] During their research, the inventors discovered that three-ring or higher aromatic hydrocarbon molecules have a higher competitive ability for adsorption in the active centers of catalysts with higher L values, generally participating in the reaction preferentially over single- and double-ring aromatic hydrocarbons. Furthermore, due to the low temperature in the hydrorefining reaction zone (I), only partial saturation of the three-ring or higher aromatic hydrocarbon molecules occurs, i.e., hydrogenation and saturation conversion to monocyclic aromatic hydrocarbon compounds with cycloalkane rings. By regulating the saturation rate of the three-ring or higher aromatic hydrocarbons, specifically by controlling the conditions of the first reaction to achieve a saturation rate of 65% to 85% for the three-ring or higher aromatic hydrocarbons in the reaction effluent (I), hydrogen consumption for fully saturated aromatic hydrocarbons can be reduced while also reducing the concentration of condensed-ring aromatic hydrocarbons, thereby suppressing catalyst deactivation due to carbon deposition.

[0045] At the same time, the inventors of the present invention also found during the research process that due to the exothermic reactions of hydrodesulfurization, denitrogenation and aromatic saturation reaction temperature occurring in the fixed-bed hydrorefining reactor, by setting two hydrorefining reaction zones with reaction temperatures from low to high, on the one hand, the fuel gas consumption of the heating furnace can be reduced, and at the same time, cold hydrogen between beds can be saved, thereby reducing the operating costs of the device; in addition, by using a hydrorefining catalyst I with a larger L value and higher activity in the hydrorefining reaction zone I at a lower temperature to convert three-ring or higher aromatic hydrocarbons into single-ring aromatic hydrocarbon compounds containing cycloalkane rings, and using a hydrorefining catalyst II with a reduced L value and activity in the hydrorefining reaction zone II at a higher temperature, and utilizing pore size adsorption diffusion limitation, it is possible to achieve selective conversion of two-ring or higher aromatic hydrocarbon compounds while retaining single-ring aromatic hydrocarbons, thereby saving reaction hydrogen consumption.

[0046] Preferably, the conditions of the first reaction are controlled so that the saturation rate of three-ring or higher aromatic hydrocarbons in the reaction effluent I is 75%-85%.

[0047] Preferably, in step (2), the conditions of the second reaction are controlled so that the saturation rate of two-ring or higher aromatic hydrocarbons in the reaction effluent II is 70%-90%, preferably 74%-90%. The inventors of the present invention have found that under this preferred condition, a higher quality tail oil fraction can be obtained and the operating cycle of the device is longer.

[0048] Preferably, in step (3), the conditions of the third reaction are controlled so that the yield of the tail oil fraction is 35%-60%. The inventors of the present invention have found that this preferred embodiment can promote the conversion of polycyclic hydrocarbons into monocyclic alkanes, monocyclic aromatic hydrocarbons or paraffins, thereby increasing the hydrogen content and UOP K value of the tail oil fraction.

[0049] Preferably, in step (1) and step (2), the loading volume ratio of the hydrorefining catalyst I to the hydrorefining catalyst II is 0.5-4:1. The inventors of the present invention have found that this preferred case is conducive to controlling the temperature gradient between the hydrorefining reaction zone I and the hydrorefining reaction zone II, can reduce the concentration of condensed ring aromatics in the reactant stream while saving hydrogen consumption in the selective saturation of aromatics, inhibit catalyst deactivation, and can also reduce the fuel gas consumption of the feed heating furnace and save interbed cooling hydrogen, thereby extending the operating cycle and saving energy and consumption.

[0050] Preferably, in step (4), the initial boiling point of the tail oil fraction is 340°C-360°C.

[0051] Preferably, in step (1), the hydrorefining catalyst I contains a carrier I and an active metal component I supported on the carrier I, the carrier I is alumina or silica-alumina, and the active metal component I contains at least one metal element selected from Group VIII non-precious metal elements and Group VIB non-precious metal elements; and the content of the active metal component I calculated as oxide is 35 wt% to 55 wt% based on the total weight of the hydrorefining catalyst I.

[0052] More preferably, in step (1), in the hydrotreating catalyst I, the Group VIII non-noble metal element is selected from at least one of nickel and cobalt, and the Group VIB non-noble metal element is a combination of molybdenum and tungsten.

[0053] Further preferably, in step (1), based on the total weight of the hydrotreating catalyst I, the content of the non-noble metal elements of Group VIII calculated as oxides is 1 wt%-15 wt%, and the content of the non-noble metal elements of Group VIB calculated as oxides is 5 wt%-40 wt%.

[0054] Preferably, in step (1), in the hydrorefining catalyst I, in the carrier I, the ratio of the total volume of pores with a pore diameter greater than or equal to 10 nm to the total pore volume is 0.6-0.8:1. The total pore volume is the total pore volume of the carrier I.

[0055] Preferably, in step (1), in the hydrorefining catalyst I, in the carrier I, the ratio of the mesopore area to the total specific surface area is 0.3-0.7:1. The mesopore area is also the surface area. The total specific surface area is the total specific surface area of the carrier I.

[0056] According to a preferred embodiment, in step (1), according to the flow direction of the liquid phase material, in the hydrorefining reaction zone I, a hydrogenation protection catalyst is loaded upstream of the hydrorefining catalyst I, and optionally, at least one catalyst selected from a hydrodemetallization catalyst and a hydroremoval of carbon residue catalyst is loaded between the hydrorefining protection catalyst and the hydrorefining catalyst I.

[0057] More preferably, in step (1), a hydrodemetallization catalyst and a hydrodecarbonization catalyst are loaded between the hydroprotection catalyst and the hydrorefining catalyst I, and the hydrodemetallization catalyst is loaded upstream of the hydrodecarbonization catalyst.

[0058] Further preferably, in step (1), based on the total volume of the catalyst in the hydrorefining reaction zone I, the loading volume content of the hydrogenation protection catalyst is 3%-20%, the loading volume content of the hydrodemetallization catalyst is 0.01%-30%, the loading volume content of the hydroremoval of carbon residue catalyst is 0.01%-30%, and the loading volume content of the hydrorefining catalyst I is 30%-96%.

[0059] Particularly preferably, in step (1), based on the total volume of the catalyst in the hydrorefining reaction zone I, the loading volume content of the hydrorefining catalyst I is 30%-60%.

[0060] Preferably, in step (1), the hydrogenation protected catalyst contains a carrier II and / or an active metal component II supported on the carrier II, the carrier II is selected from at least one of aluminum oxide, silicon oxide and titanium oxide, and the active metal component II contains at least one metal element selected from Group VIII non-precious metal elements and Group VIB metal elements; and the content of the active metal component II calculated as oxide is 0.1 wt% to 15 wt% based on the total weight of the hydrogenation protected catalyst.

[0061] Preferably, the linear average particle size of the hydrogenation protection catalyst is 0.5 mm to 50.0 mm, and the bulk density is 0.3 g / cm 3 -1.2g / cm 3 , with a specific surface area of 50m 2 / g-300m 2 / g.

[0062] Preferably, in step (1), the hydrodemetallization catalyst contains a carrier III and an active metal component III supported on the carrier III, the carrier III is selected from at least one of aluminum oxide, silicon oxide and titanium oxide, and the active metal component III contains at least one metal element selected from Group VIII non-noble metal elements and Group VIB metal elements; and the content of the active metal component III calculated as oxide is 3 wt% to 30 wt% based on the total weight of the hydrodemetallization catalyst.

[0063] Preferably, the linear average particle size of the hydrodemetallization catalyst is 0.2 mm to 6.0 mm, and the bulk density is 0.3 g / cm 3 -0.8g / cm 3 , with a specific surface area of 100m 2 / g-250m 2 / g.

[0064] Preferably, in step (1), the hydrodecarbonization catalyst contains a carrier IV and an active metal component IV supported on the carrier IV, the carrier IV is selected from at least one of aluminum oxide, silicon oxide and titanium oxide, and the active metal component IV contains at least one metal element selected from Group VIII non-noble metal elements and Group VIB metal elements; and the content of the active metal component IV in terms of oxide is 3 wt% to 33 wt% based on the total weight of the hydrodecarbonization catalyst.

[0065] Preferably, the linear average particle size of the hydroremoval of carbon residue catalyst is 0.2 mm to 3.0 mm, and the bulk density is 0.3 g / cm 3 -0.7g / cm 3 , with a specific surface area of 100m 2 / g-250m 2 / g.

[0066] According to another preferred embodiment, in step (2), the hydrorefining catalyst II contains a carrier V and an active metal component V supported on the carrier V, the carrier V is alumina or silica-alumina, and the active metal component V contains at least one metal element selected from Group VIII non-precious metal elements and Group VIB non-precious metal elements; and based on the total weight of the hydrorefining catalyst II, the content of the Group VIII non-precious metal element is 1 wt%-15 wt%, and the content of the Group VIB non-precious metal element calculated as oxide is 5 wt%-40 wt%.

[0067] More preferably, in step (2), in the hydrotreating catalyst II, the Group VIII non-noble metal element is selected from at least one of nickel and cobalt, and the Group VIB non-noble metal element is selected from at least one of molybdenum and tungsten.

[0068] Furthermore, in step (2), in the hydrorefining catalyst II, in the carrier V, the ratio of the total volume of pores with a pore diameter greater than or equal to 10 nm to the total pore volume is 0.4-0.59:1. The total pore volume is the total pore volume of the carrier V.

[0069] According to a particularly preferred embodiment, the ratio of the total volume of pores with a pore diameter greater than or equal to 10 nm to the total pore volume is defined as L, and the L value of the carrier I is greater than the L value of the carrier V. The inventors of the present invention have found that under this preferred embodiment, the hydrotreating catalyst can have higher activity.

[0070] Preferably, in step (3), the hydrocracking catalyst contains a carrier VI and an active metal component VI supported on the carrier VI, the carrier VI contains a heat-resistant inorganic oxide and an acidic component, the heat-resistant inorganic oxide is selected from at least one of silicon oxide and aluminum oxide, and the acidic component contains Y-type molecular sieve and amorphous silicon aluminum; the active metal component VI contains at least two metal elements selected from Group VIB metal elements and Group VIII metal elements; and based on the total weight of the hydrocracking catalyst, the content of the Group VIB metal element calculated as oxide is 15wt%-33wt%, and / or the content of the Group VIII metal element calculated as oxide is 2wt%-8wt%.

[0071] Further preferably, based on the total weight of the carrier VI, the content of the acidic component is 10 wt%-65 wt%.

[0072] Particularly preferably, based on the total weight of the acidic components, the content of the Y-type molecular sieve is 60 wt%-100 wt%.

[0073] According to a preferred embodiment, in step (1) and step (2), the conditions of the first reaction and the second reaction independently include: hydrogen partial pressure of 6.0MPa-20.0MPa, reaction temperature of 280℃-400℃, liquid hourly space velocity of 0.5h -1 -6h -1 , the hydrogen-to-oil volume ratio is 300-2000.

[0074] According to another preferred embodiment, the conditions of the third reaction include: in step (3), the hydrogen partial pressure is 6.0MPa-20.0MPa, the reaction temperature is 290℃-420℃, the liquid hourly space velocity is 0.3h -1 -5h -1 , the hydrogen-to-oil volume ratio is 300-2000.

[0075] Preferably, the heavy low-quality wax oil is selected from at least one of straight-run wax oil and secondary processed oil.

[0076] More preferably, the straight-run wax oil is selected from at least one of second-line wax oil, third-line wax oil, fourth-line wax oil and fifth-line wax oil; the secondary processing oil is selected from at least one of slurry bed wax oil, boiling bed wax oil and deasphalted oil.

[0077] The present invention will be described in detail below through examples.

[0078] In the following examples, unless otherwise specified, all raw materials used were commercially available.

[0079] In the following examples, the preparation method of purified CAT1 is as follows:

[0080] Weigh 1000g of pseudo-boehmite (purchased from Sinopec Catalyst Changling Branch, with a dry basis of 0.73 and a specific surface area of 300m 2 / g, pore volume of 0.97 mL / g), extruded into clover-shaped strips with a circumscribed circle diameter of 1.6 mm, dried at 110 ° C for 8 h, ventilated with air, heated to 600 ° C at a rate of 3 ° C / min, and calcined for 3 h to prepare alumina support I1;

[0081] 200 g of carrier I1 was placed in a polytetrafluoroethylene-lined self-pressurized sealed autoclave, 100 g of deionized water was added and sealed, and the autoclave was subjected to a constant temperature hydrothermal treatment at 90°C for 12 h, followed by filtration. The obtained solid was dried at 120°C for 4 h to obtain carrier I2;

[0082] At room temperature (25±2°C), 200g of carrier I2 was placed in 242mL of an aqueous solution containing 68.5g of ammonium metatungstate, 13.7g of ammonium paramolybdate, 20.1g of basic nickel carbonate and 3.2g of boric acid, and the mixture was immersed for 1h and filtered. The obtained solid was dried at 120°C for 6h and then calcined at 400°C for 3h to obtain carrier I3; carrier I3 was placed in 242mL of an aqueous solution containing 23.14g of citric acid and immersed for 1h and filtered. The obtained solid was dried at 120°C for 3h to obtain refined CAT1.

[0083] In the following example,

[0084] The yield of the light naphtha fraction is defined as: the weight of the separated light naphtha fraction (<65°C) / the weight of the heavy inferior wax oil × 100%;

[0085] The yield of the heavy naphtha fraction is defined as: the weight of the separated heavy naphtha fraction (65°C-175°C) / the weight of the heavy inferior wax oil × 100%;

[0086] The yield of the diesel fraction is defined as: the weight of the separated middle distillate oil (175°C - T1°C) / the weight of the heavy inferior wax oil × 100%; T1 is the initial boiling point of the tail oil fraction.

[0087] raw material:

[0088] Hydrorefining catalyst:

[0089] Refined CAT1: Carrier I is alumina, and as oxides, the content of molybdenum is 5.6wt%, the content of tungsten is 32wt%, and the content of nickel is 6.3wt%. The L value of carrier I is 0.693:1, and the surface area of carrier I is 71.68m 2 / g, and the ratio of the total specific surface area is 0.3106:1;

[0090] Refined CAT2-1: The brand is RN-32V, the carrier V is alumina, and the tungsten content, calculated as oxide, is 23.0wt%, the molybdenum content is 2.3wt%, and the nickel content is 2.4wt%. The L value of the carrier V is 0.5675:1;

[0091] Hydrogenation protection catalyst:

[0092] Protective CAT3: Brand is RG-200, carrier II is alumina, does not contain active metal components, linear average particle size is 16mm, specific surface area is 100m 2 / g, bulk density is 0.65g / cm 3 ;

[0093] Protective CAT4: Brand is RG-201, carrier II is alumina, calculated as oxide, the content of molybdenum is 1.2wt%, the content of nickel is 0.3wt%, the linear average particle size is 10mm, and the specific surface area is 95m 2 / g, bulk density is 0.56g / cm 3 ;

[0094] Protection CAT5: Brand is RG-30A, carrier II is alumina, calculated as oxide, the molybdenum content is 2.5wt%, the nickel content is 0.5wt%, the linear average particle size is 6.0mm, and the specific surface area is 90m 2 / g, bulk density is 0.48g / cm 3 ;

[0095] Protection CAT6: Brand is RG-30B, carrier II is alumina, calculated as oxide, the molybdenum content is 5.5wt%, the nickel content is 1.0wt%, the linear average particle size is 3.0mm, and the specific surface area is 90.0m 2 / g, bulk density is 0.50g / cm 3 ;

[0096] Hydrodemetallization catalyst:

[0097] Demetallized CAT7: The brand is RAM-100, the carrier III is alumina, the content of molybdenum is 7.5wt%, the content of nickel is 1.3wt%, the linear average particle size is 1.8mm, and the specific surface area is 160m 2 / g, bulk density is 0.45g / cm 3 ;

[0098] Hydrodecarbonization catalyst:

[0099] Residual carbon removal CAT8: The brand is RCS-31, the carrier IV is alumina, the molybdenum content is 15.8wt% and the nickel content is 3.2wt% in terms of oxide; the linear average particle size is 1.3mm and the specific surface area is 160m 2 / g, bulk density is 0.61g / cm 3 ;

[0100] Hydrocracking catalyst:

[0101] Cracking CAT9: RHC-210, with a tungsten content of 27.0 wt% and a nickel content of 2.7 wt% based on oxides, with the balance being carrier VI; an acidic component content of 65 wt% based on the total weight of carrier VI, with the balance being alumina; and a Y-type molecular sieve content of 100 wt% based on the total weight of the acidic components;

[0102] Cracking CAT10: The brand is RHC-220. Calculated as oxides, the tungsten content is 27.0wt%, the nickel content is 2.7wt%, and the balance is carrier VI; based on the total weight of carrier VI, the content of the acidic component is 25wt%, and the balance is alumina; based on the total weight of the acidic components, the content of Y-type molecular sieve is 100wt%.

[0103] In the following examples, the properties of the heavy low-quality wax oil used are shown in Table 1.

[0104] Table 1

[0105]

[0106]

[0107] Example 1

[0108] (1) introducing the heavy low-quality wax oil into the hydrotreating reaction zone I filled with the hydrotreating catalyst I to perform a first reaction to obtain a reaction effluent I;

[0109] (2) introducing the reaction effluent I into the hydrotreating reaction zone II filled with a hydrotreating catalyst II to perform a second reaction to obtain a reaction effluent II;

[0110] (3) introducing the reaction effluent II into the hydrocracking reaction zone filled with a hydrocracking catalyst to perform a third reaction to obtain a reaction effluent III;

[0111] (4) separating the reaction effluent III to obtain a light naphtha fraction, a heavy naphtha fraction, a middle distillate oil, and a tail oil fraction;

[0112] The specific process parameters are shown in Table 2, and the yield and property parameters of the product are shown in Table 3.

[0113] Example 2

[0114] This embodiment is carried out in the same manner as in embodiment 1, except that:

[0115] The reaction temperature for the first reaction was 375°C;

[0116] The reaction temperature of the third reaction is 370°C;

[0117] The remaining process parameters are the same as those in Example 1.

[0118] In this embodiment, the saturation rate of three-ring or higher aromatic hydrocarbons in the reaction effluent I is 83.46%;

[0119] The nitrogen mass fraction in the reaction effluent II is 10 μg·g -1 , the saturation rate of two-ring and above aromatic hydrocarbons is 87.21%;

[0120] The yield of tail oil fraction is 42.2%.

[0121] The yield and property parameters of the product are shown in Table 3.

[0122] Example 3

[0123] This embodiment is carried out in the same manner as in embodiment 1, except that:

[0124] The reaction temperature for the first reaction was 375°C;

[0125] The reaction temperature of the third reaction is 375°C;

[0126] The remaining process parameters are the same as those in Example 1.

[0127] In this embodiment, the saturation rate of three-ring or higher aromatic hydrocarbons in the reaction effluent I is 83.46%;

[0128] The nitrogen mass fraction in the reaction effluent II is 10 μg·g -1 , the saturation rate of two-ring and above aromatic hydrocarbons is 87.21%;

[0129] The yield of tail oil fraction is 35.0%.

[0130] The yield and property parameters of the product are shown in Table 3.

[0131] Example 4

[0132] This embodiment is carried out in the same manner as in embodiment 1, except that:

[0133] The reaction temperature for the first reaction was 375°C;

[0134] The reaction temperature of the third reaction is 364°C;

[0135] The remaining process parameters are the same as those in Example 1.

[0136] In this embodiment, the saturation rate of three-ring or higher aromatic hydrocarbons in the reaction effluent I is 83.46%;

[0137] The nitrogen mass fraction in the reaction effluent II is 10 μg·g -1 , the saturation rate of two-ring and above aromatic hydrocarbons is 87.21%;

[0138] The yield of tail oil fraction is 50.0%.

[0139] The yield and property parameters of the product are shown in Table 3.

[0140] Example 5

[0141] This embodiment is carried out in the same manner as in embodiment 1, except that:

[0142] The reaction temperature of the first reaction is 380°C;

[0143] The loading volume ratio of hydrofining catalyst I to hydrofining catalyst II is 1:1;

[0144] The reaction temperature of the third reaction is 368°C;

[0145] The remaining process parameters are the same as those in Example 1.

[0146] In this embodiment, the saturation rate of three-ring or higher aromatic hydrocarbons in the reaction effluent I is 80.46%;

[0147] The nitrogen mass fraction in the reaction effluent II is 16 μg·g -1 , the saturation rate of two-ring and above aromatic hydrocarbons is 85.21%;

[0148] The yield of tail oil fraction is 45.0%.

[0149] The yield and property parameters of the product are shown in Table 3.

[0150] Example 6

[0151] This embodiment is carried out in the same manner as in embodiment 1, except that:

[0152] The reaction temperature of the first reaction is 380°C;

[0153] The loading volume ratio of hydrofining catalyst I to hydrofining catalyst II is 1:1;

[0154] The reaction temperature of the third reaction is 373°C;

[0155] The remaining process parameters are the same as those in Example 1.

[0156] In this embodiment, the saturation rate of three-ring or higher aromatic hydrocarbons in the reaction effluent I is 80.46%;

[0157] The nitrogen mass fraction in the reaction effluent II is 16 μg·g -1 , the saturation rate of two-ring and above aromatic hydrocarbons is 85.21%;

[0158] The yield of tail oil fraction is 38.0%.

[0159] The yield and property parameters of the product are shown in Table 3.

[0160] Example 7

[0161] This embodiment is carried out in the same manner as in embodiment 1, except that:

[0162] The reaction temperature of the first reaction is 380°C;

[0163] The loading volume ratio of hydrofining catalyst I to hydrofining catalyst II is 1:1;

[0164] The reaction temperature of the third reaction is 385°C;

[0165] In the hydrocracking reaction zone, an equal volume of cracking CAT9 is replaced by cracking CAT10;

[0166] The remaining process parameters are the same as those in Example 1.

[0167] In this embodiment, the saturation rate of three-ring or higher aromatic hydrocarbons in the reaction effluent I is 80.46%;

[0168] The nitrogen mass fraction in the reaction effluent II is 16 μg·g -1 , the saturation rate of two-ring and above aromatic hydrocarbons is 85.21%;

[0169] The yield of tail oil fraction is 39.0%.

[0170] The yield and property parameters of the product are shown in Table 3.

[0171] Example 8

[0172] This embodiment is carried out in the same manner as in embodiment 1, except that:

[0173] The reaction temperature of the first reaction was 384°C;

[0174] The loading volume ratio of hydrofining catalyst I to hydrofining catalyst II is 1:1;

[0175] The reaction temperature of the second reaction is 399°C;

[0176] The reaction temperature of the third reaction is 360°C;

[0177] The remaining process parameters are the same as those in Example 1.

[0178] In this embodiment, the saturation rate of three-ring or higher aromatic hydrocarbons in the reaction effluent I is 84.21%;

[0179] The nitrogen mass fraction in the reaction effluent II is 5 μg·g -1 , the saturation rate of two-ring and above aromatic hydrocarbons is 91.21%;

[0180] The yield of tail oil fraction is 55.0%.

[0181] The yield and property parameters of the product are shown in Table 3.

[0182] Example 9

[0183] This embodiment is carried out in the same manner as in embodiment 1, except that:

[0184] The reaction temperature for the first reaction was 365°C;

[0185] In the hydrofining reaction zone I, an equal volume of refined CAT1 was replaced with refined CAT2-1;

[0186] In the hydrofining reaction zone II, an equal volume of refined CAT2-1 was replaced with refined CAT1;

[0187] The reaction temperature of the third reaction is 374°C;

[0188] The remaining process parameters are the same as those in Example 1.

[0189] In this embodiment, the saturation rate of three-ring or higher aromatic hydrocarbons in the reaction effluent I is 71.63%;

[0190] The nitrogen mass fraction in the reaction effluent II is 54 μg·g -1 , the saturation rate of two-ring and above aromatic hydrocarbons is 71.35%;

[0191] The yield of tail oil fraction is 47.5%.

[0192] The yield and property parameters of the product are shown in Table 3.

[0193] Example 10

[0194] This embodiment is carried out in the same manner as in embodiment 1, except that:

[0195] The reaction temperature of the third reaction is 354°C;

[0196] The remaining process parameters are the same as those in Example 1.

[0197] In this embodiment, the tail oil fraction yield is 64.0%.

[0198] The yield and property parameters of the product are shown in Table 3.

[0199] Example 11

[0200] This embodiment is carried out in the same manner as in embodiment 1, except that:

[0201] The reaction temperature of the first reaction was 390°C;

[0202] The loading volume ratio of hydrofining catalyst I to hydrofining catalyst II is 1:3;

[0203] The reaction temperature of the third reaction is 370°C;

[0204] The remaining process parameters are the same as those in Example 1.

[0205] In this embodiment, the saturation rate of three-ring or higher aromatic hydrocarbons in the reaction effluent I is 65.10%;

[0206] The nitrogen mass fraction in the reaction effluent II is 60 μg·g -1 , the saturation rate of two-ring or higher aromatic hydrocarbons is 70.10%;

[0207] The yield of tail oil fraction is 52.0%.

[0208] The yield and property parameters of the product are shown in Table 3.

[0209] Comparative Example 1

[0210] This comparative example was carried out in the same manner as in Example 1, except that:

[0211] The reaction temperature of the first reaction is 350°C;

[0212] The reaction temperature of the third reaction is 373°C;

[0213] The remaining process parameters are the same as those in Example 1.

[0214] In this embodiment, the saturation rate of three-ring or higher aromatic hydrocarbons in the reaction effluent I is 61.92%;

[0215] The nitrogen mass fraction in the reaction effluent II is 51 μg·g -1 , the saturation rate of two-ring and above aromatic hydrocarbons is 72.40%;

[0216] The yield of tail oil fraction is 47.0%.

[0217] The yield and property parameters of the product are shown in Table 3.

[0218] Comparative Example 2

[0219] This comparative example was carried out in the same manner as in Example 1, except that:

[0220] The reaction temperature of the first reaction was 390°C;

[0221] The reaction temperature of the second reaction was 392°C;

[0222] The reaction temperature of the third reaction is 360°C;

[0223] The remaining process parameters are the same as those in Example 1.

[0224] In this embodiment, the saturation rate of three-ring or higher aromatic hydrocarbons in the reaction effluent I is 87.10%;

[0225] The nitrogen mass fraction in the reaction effluent II is 7 μg·g -1 , the saturation rate of two-ring and above aromatic hydrocarbons is 88.31%;

[0226] The yield of tail oil fraction is 55.6%.

[0227] The yield and property parameters of the product are shown in Table 3.

[0228] Table 2

[0229]

[0230] Note: 1. Protection CAT3+Protection CAT4+Protection CAT5+Protection CAT6+Demetallization CAT7+Residual Carbon Removal CAT8+Refined CAT1 means that according to the direction of liquid material flow, Protection CAT3, Protection CAT4, Protection CAT5, Protection CAT6, Demetallization CAT7, Residual Carbon Removal CAT8 and Refined CAT1 are loaded in sequence;

[0231] 2. The column of "Packing Volume Ratio" refers to the loading volume ratio of hydrotreating catalyst I to hydrotreating catalyst II.

[0232] Table 3

[0233]

[0234] From the data in Table 2 and Table 3, we can see that:

[0235] In Examples 1 to 7, the saturation rate of three-ring or higher aromatic hydrocarbons in the reaction effluent I was controlled within the range of 65%-85%; and the reaction temperature in the hydrotreating reaction zone I was controlled to be 10°C-30°C lower than the reaction temperature in the hydrotreating reaction zone II. The resulting tail oil UPO K value was not less than 12.0, and the tail oil hydrogen content was not less than 13.5 wt%. It can be seen that the method provided by the present invention can meet the requirements for high-quality DCC unit raw material feed.

[0236] In Example 8, by significantly increasing the reaction temperature of the hydrorefining reaction zone I and the hydrorefining reaction zone II, the saturation rate of two-ring and higher aromatic hydrocarbons in the reaction effluent II was 91.21%, and the heavy low-quality wax oil was ultra-deeply hydrogenated and saturated, and the product quality was correspondingly improved. However, the operating temperature of the hydrorefining reaction zone I and the hydrorefining reaction zone II was relatively high, which shortened the operating cycle of the device and was not conducive to long-term operation.

[0237] In Example 9, the hydrorefining catalyst CAT2-1 with a lower L value was loaded into the hydrorefining reaction zone I, and the hydrorefining catalyst CAT1 with a higher L value was loaded into the hydrorefining reaction zone II. Under similar reaction conditions, due to the lower L value of the hydrorefining catalyst CAT2-1, the number of active centers in the catalyst that could be contacted by the heavy, low-quality wax oil macromolecular aromatic compounds was small, resulting in a low saturation rate of three-ring or larger aromatics in the reaction effluent I, which in turn led to a lower saturation rate of two-ring or larger aromatics in the hydrorefining reaction zone II and an increase in the nitrogen content of the refined oil. Although reaction compensation was achieved by increasing the reaction temperature in the hydrocracking reaction zone, the excessively high nitrogen content of the refined oil could lead to poisoning of the hydrocracking catalyst and accelerated catalyst deactivation, which is not conducive to long-term operation.

[0238] In Example 10, the reaction temperature in the hydrocracking reaction zone was 354°C, and when the yield of the tail oil fraction was 64.0%, the hydrogen content of the tail oil was less than 13.5 wt%, and the UOP K value of the tail oil was less than 12.0, which could not meet the quality requirements of the high-quality DCC unit feedstock.

[0239] In Example 11, reducing the loading volume ratio of the hydrorefining catalyst in the hydrorefining reaction zone I will, on the one hand, lead to an excessively high operating temperature of the hydrorefining reaction zone I, resulting in an increase in the energy consumption of the heating furnace, which is not conducive to reducing the operating cost of the device; in addition, the loading volume ratio of the catalyst in the hydrorefining reaction zone I is reduced, and the saturation rate of three-ring or larger aromatic hydrocarbons in the reaction effluent I is reduced, which is not conducive to the removal of nitrogen-containing compounds and the long cycle life of the hydrorefining catalyst in the hydrorefining reaction zone; while the reaction temperature in the hydrorefining reaction zone II is controlled unchanged, the saturation rate of two-ring or larger aromatic hydrocarbons in the reaction effluent II is 70.10%, and the nitrogen content is 60.0 μg / g. It is necessary to increase the reaction temperature of the hydrocracking reaction zone to compensate for the effects of denitrogenation and insufficient saturation depth of two-ring or larger aromatic hydrocarbons.

[0240] In Comparative Example 1, the reaction temperature of the hydrorefining reaction zone I is 350° C., the saturation rate of three-ring or larger aromatics in the reaction effluent I is 61.92%, the reaction temperature of the hydrorefining reaction zone II is 390° C., and the nitrogen content in the reaction effluent II is 51.0 μg / g. Due to the large temperature difference between the hydrorefining reaction zone I and the hydrorefining reaction zone II, the saturation depth of three-ring or larger aromatics in the reaction effluent I is insufficient, which increases the difficulty of denitrogenation and aromatic saturation in the hydrorefining reaction zone II. On the one hand, it is necessary to increase the hydrocracking reaction temperature to compensate for the effects of insufficient denitrogenation and three-ring or larger aromatic saturation depth. On the other hand, a higher nitrogen content will accelerate the deactivation rate of the hydrocracking catalyst, which is not conducive to the long-term operation of the device.

[0241] In Comparative Example 2, the reaction temperature of the hydrorefining reaction zone I was increased to a level close to that of the hydrorefining reaction zone II, so that the saturation rate of three-ring or larger aromatic hydrocarbons in the reaction effluent I was 87.10%, and the saturation rate of two-ring or larger aromatic hydrocarbons in the reaction effluent II was 88.31%. Although qualified products were obtained, the hydrorefining reaction zone I and the hydrorefining reaction zone II adopted similar operating temperatures. This, on the one hand, would lead to a substantial increase in the fuel gas consumption of the inlet heating furnace of the hydrorefining reaction zone I, and a large amount of cold hydrogen would need to be used to control the catalyst bed temperature, which would substantially increase the operating costs and energy consumption. In addition, a substantial increase in the operating severity of the hydrorefining reaction zone I would lead to an accelerated deactivation rate of the hydrorefining catalyst, which would be detrimental to the long-term operation of the device.

[0242] The preferred embodiments of the present invention have been described in detail above, but the present invention is not limited thereto. Within the technical concept of the present invention, various simple variations of the technical solution of the present invention may be made, including combining the various technical features in any other appropriate manner. These simple variations and combinations should also be regarded as disclosed in the present invention and fall within the scope of protection of the present invention.

Claims

1. A hydrocracking method for heavy low-quality wax oil, characterized in that: The method is carried out in a device containing a hydrofining reaction zone and a hydrocracking reaction zone, wherein the hydrofining reaction zone contains a hydrofining reaction zone I and a hydrofining reaction zone II. The method comprises: (1) introducing the heavy low-quality wax oil into the hydrorefining reaction zone I filled with the hydrorefining catalyst I to carry out a first reaction to obtain a reaction effluent I; (2) introducing the reaction effluent I into the hydrotreating reaction zone II filled with a hydrotreating catalyst II to carry out a second reaction to obtain a reaction effluent II; controlling the conditions of the second reaction so that the saturation rate of two-ring or higher aromatic hydrocarbons in the reaction effluent II is 70%-90%; (3) introducing the reaction effluent II into the hydrocracking reaction zone filled with a hydrocracking catalyst to perform a third reaction to obtain a reaction effluent III; (4) Separating the reaction effluent III to obtain a light naphtha fraction, a heavy naphtha fraction, an intermediate distillate oil, and a tail oil fraction; controlling the conditions of the third reaction so that the yield of the tail oil fraction is 35%-60%; and the initial boiling point of the tail oil fraction is 340°C-360°C; The heavy low-quality wax oil has a 95% distillation temperature of 580° C. to 700° C. according to ASTM D-1160, a nitrogen content of 1800 μg / g to 3500 μg / g, and a sum of two-ring or higher cycloalkanes and two-ring or higher aromatics of 50 wt % to 80 wt %. Controlling the conditions of the first reaction so that the saturation rate of three-ring or higher aromatic hydrocarbons in the reaction effluent I is 65%-85%; The reaction temperature of the first reaction is 10°C-30°C lower than the reaction temperature of the second reaction; In step (1) and step (2), the loading volume ratio of the hydrotreating catalyst I to the hydrotreating catalyst II is 0.5-4:

1.

2. The method according to claim 1, wherein In step (1), the hydrorefining catalyst I contains a carrier I and an active metal component I supported on the carrier I, the carrier I is alumina or silica-alumina, and the active metal component I contains at least one metal element selected from Group VIII non-precious metal elements and Group VIB non-precious metal elements; and the content of the active metal component I calculated as oxide is 35 wt% to 55 wt% based on the total weight of the hydrorefining catalyst I.

3. The method according to claim 2, wherein: In step (1), in the hydrorefining catalyst I, the Group VIII non-noble metal element is selected from at least one of nickel and cobalt, and the Group VIB non-noble metal element is a combination of molybdenum and tungsten.

4. The method according to claim 3, wherein: In step (1), based on the total weight of the hydrorefining catalyst I, the content of the non-noble metal elements of Group VIII calculated as oxides is 1 wt%-15 wt%, and the content of the non-noble metal elements of Group VIB calculated as oxides is 5 wt%-40 wt%.

5. The method according to claim 2, wherein: In step (1), in the hydrorefining catalyst I, in the carrier I, the ratio of the total volume of pores with a pore diameter greater than or equal to 10 nm to the total pore volume is 0.6-0.8:

1.

6. The method according to claim 5, wherein: In step (1), in the hydrorefining catalyst I, in the carrier I, the ratio of the mesopore area to the total specific surface area is 0.3-0.7:

1.

7. The method according to any one of claims 1 to 6, wherein: In step (1), according to the flow direction of the liquid phase material, in the hydrorefining reaction zone I, a hydrogenation protection catalyst is loaded upstream of the hydrorefining catalyst I, and optionally, at least one catalyst selected from the group consisting of a hydrodemetallization catalyst and a hydroremoval of carbon residue catalyst is loaded between the hydrorefining protection catalyst and the hydrorefining catalyst I.

8. The method according to claim 7, wherein: In step (1), a hydrodemetallization catalyst and a hydrodecarbonization catalyst are loaded between the hydroprotection catalyst and the hydrorefining catalyst I, and the hydrodemetallization catalyst is loaded upstream of the hydrodecarbonization catalyst.

9. The method according to claim 8, wherein In step (1), based on the total volume of the catalyst in the hydrorefining reaction zone I, the loading volume content of the hydrogenation protection catalyst is 3%-20%, the loading volume content of the hydrodemetallization catalyst is 0.01%-30%, the loading volume content of the hydroremoval of carbon residue catalyst is 0.01%-30%, and the loading volume content of the hydrorefining catalyst I is 30%-96%.

10. The method according to claim 7, wherein: In step (1), the hydrogenation protected catalyst contains a carrier II and an active metal component II supported on the carrier II, the carrier II is selected from at least one of aluminum oxide, silicon oxide and titanium oxide, and the active metal component II contains at least one metal element selected from Group VIII non-noble metal elements and Group VIB metal elements; and the content of the active metal component II calculated as oxide is 0.1 wt% to 15 wt% based on the total weight of the hydrogenation protected catalyst.

11. The method according to claim 10, wherein: The linear average particle size of the hydrogenation protection catalyst is 0.5 mm to 50.0 mm, and the bulk density is 0.3 g / cm 3 -1.2g / cm 3 , with a specific surface area of 50m 2 / g-300m 2 / g.

12. The method according to claim 7, wherein: In step (1), the hydrodemetallization catalyst contains a carrier III and an active metal component III supported on the carrier III, the carrier III is selected from at least one of aluminum oxide, silicon oxide and titanium oxide, and the active metal component III contains at least one metal element selected from Group VIII non-noble metal elements and Group VIB metal elements; and the content of the active metal component III calculated as oxide is 3 wt% to 30 wt% based on the total weight of the hydrodemetallization catalyst.

13. The method according to claim 12, wherein: The linear average particle size of the hydrodemetallization catalyst is 0.2 mm to 6.0 mm, and the bulk density is 0.3 g / cm 3 -0.8g / cm 3 , with a specific surface area of 100m 2 / g-250m 2 / g.

14. The method according to claim 7, wherein: In step (1), the hydrodecarbonization catalyst contains a carrier IV and an active metal component IV supported on the carrier IV, the carrier IV is selected from at least one of aluminum oxide, silicon oxide and titanium oxide, and the active metal component IV contains at least one metal element selected from Group VIII non-noble metal elements and Group VIB metal elements; and the content of the active metal component IV in terms of oxide is 3 wt% to 33 wt% based on the total weight of the hydrodecarbonization catalyst.

15. The method according to claim 14, wherein The linear average particle size of the hydrogenation carbon residue removal catalyst is 0.2 mm to 3.0 mm, and the bulk density is 0.3 g / cm 3 -0.7g / cm 3 , with a specific surface area of 100m 2 / g-250m 2 / g.

16. The method according to any one of claims 1 to 6, wherein: In step (2), the hydrorefining catalyst II contains a carrier V and an active metal component V supported on the carrier V, wherein the carrier V is alumina or silica-alumina, and the active metal component V contains at least one metal element selected from Group VIII non-precious metal elements and Group VIB non-precious metal elements.

17. The method according to claim 16, wherein In step (2), based on the total weight of the hydrotreating catalyst II, the content of the non-noble metal elements of Group VIII is 1 wt%-15 wt%, and the content of the non-noble metal elements of Group VIB in terms of oxide is 5 wt%-40 wt%.

18. The method according to claim 17, wherein In step (2), in the hydrorefining catalyst II, the Group VIII non-noble metal element is selected from at least one of nickel and cobalt, and the Group VIB non-noble metal element is selected from at least one of molybdenum and tungsten.

19. The method according to claim 18, wherein In step (2), in the hydrorefining catalyst II, in the carrier V, the ratio of the total volume of pores with a pore diameter greater than or equal to 10 nm to the total pore volume is 0.4-0.59:

1.

20. The method according to claim 16, wherein The ratio of the total volume of pores with a pore diameter greater than or equal to 10 nm to the total pore volume is defined as L, and the L value of the carrier I is greater than the L value of the carrier V.

21. The method according to any one of claims 1 to 6, wherein: In step (3), the hydrocracking catalyst contains a carrier VI and an active metal component VI supported on the carrier VI, the carrier VI contains a heat-resistant inorganic oxide and an acidic component, the heat-resistant inorganic oxide is selected from at least one of silicon oxide and aluminum oxide, and the acidic component contains at least one of Y-type molecular sieve and amorphous silicon aluminum; the active metal component VI contains at least two metal elements selected from Group VIB metal elements and Group VIII metal elements; and based on the total weight of the hydrocracking catalyst, the content of the Group VIB metal element calculated as oxide is 15wt%-33wt%, and / or the content of the Group VIII metal element calculated as oxide is 2wt%-8wt%.

22. The method according to claim 21, wherein Based on the total weight of the carrier VI, the content of the acidic component is 10 wt % to 65 wt %.

23. The method according to claim 21, wherein Based on the total weight of the acidic components, the content of the Y-type molecular sieve is 60wt%-100wt%.

24. The method according to any one of claims 1 to 6, wherein: In step (1) and step (2), the conditions of the first reaction and the second reaction independently include: hydrogen partial pressure of 6.0 MPa-20.0 MPa, reaction temperature of 280°C-400°C, liquid hourly space velocity of 0.5 h -1 -6h -1 , the hydrogen-to-oil volume ratio is 300-2000.

25. The method according to any one of claims 1 to 6, wherein: In step (3), the conditions of the third reaction include: hydrogen partial pressure of 6.0MPa-20.0MPa, reaction temperature of 290℃-420℃, liquid hourly space velocity of 0.3h -1 -5h -1 , the hydrogen-to-oil volume ratio is 300-2000.

Images