A method for producing high viscosity index lubricating oil base oil by hydrogenation

Abstract

The present invention discloses a method for producing high viscosity index lubricating oil base oil by hydrogenation. Waxy oil is used as raw material, and a hydrogenation treatment reaction is first performed; the effluent of the hydrogenation treatment reaction is stripped by high-pressure hydrogen to remove H that has an impact on the precious metal catalyst. 2 S and NH 3 After removing impurities, it enters the isomerization decondensation-deep supplementary refining reaction zone to improve the low temperature performance and stability of the oil. The method of the present invention can obtain a series of high viscosity index lubricating oil base oils, while producing high-quality special oil products as by-products, and greatly reduces the energy consumption caused by the traditional two-stage pressure increase-pressure reduction-pressure increase, temperature increase-temperature reduction-temperature increase, and has the advantages of simple process and small footprint.

Description

Technical Field

[0001] The invention belongs to the technical field of hydrogenation, and in particular relates to a method for producing a high viscosity index lubricating oil base oil by hydrogenation. Background Art

[0002] Traditional lubricating oil base oil production adopts solvent refining process, and its main two steps are solvent refining to remove non-ideal components such as aromatics and solvent dewaxing to ensure the low temperature flow properties of base oil. In addition, clay or hydrogenation refining is generally required.

[0003] The increasingly stringent environmental regulations and the rapid development of the machinery industry have put forward higher and higher requirements on the performance of lubricant base oils. At the same time, due to the deterioration of crude oil worldwide, the amount of crude oil suitable for the production of lubricants has gradually decreased. Therefore, the technology of producing lubricants by hydrogenation has developed rapidly. The hydrogenation process refers to the process of producing lubricant base oils by hydrocracking process or hydrotreating-isomerization dewaxing-hydrofining combined process. Its advantages are large raw material flexibility, high base oil yield, and high by-product value.

[0004] One disadvantage of the traditional hydroisomerization dewaxing process is that when full-fraction or wide-fraction waxy oil is used as feed, it is difficult to make both the light lubricating oil component and the heavy lubricating oil component meet the requirements of pour point and viscosity index at the same time. Generally speaking, when the pour point of the heavy lubricating oil component is qualified, the viscosity index loss of the light lubricating oil component is large, and it is difficult to produce API III light base oil products with a viscosity index>120; and when the viscosity index of the light lubricating oil component is qualified, the heavy component cannot be used as a qualified lubricating oil component.

[0005] In view of the above problems, the existing technology currently adopts pre-fractionation of waxy oil, and then uses each narrow fraction as feed for hydroisomerization dewaxing to solve the problem of producing high viscosity index light lubricating oil base oil. US5,580,442 proposes a method for producing high viscosity index lubricating oil base oil from hydrocracking tail oil. First, the hydrocracking tail oil is pre-fractionated under reduced pressure to cut out 100 viscosity 3mm 2 / s, 4mm 2 / s, 6 mm 2 / s and 8mm 2 / s four narrow fractions, these narrow fractions were hydroisomerized and dewaxed to produce 100 and 3mm viscosity with viscosity index>130 2 / s, 4mm 2 / s, 6 mm 2 / s and 8mm 2 / s of four lubricant base oil products.

[0006] Using narrow-fraction waxy oil as feed for hydroisomerization dewaxing can solve the problem of producing light and heavy high-viscosity index lubricant base oils, but it requires setting up multiple raw material tanks, increasing the construction investment of the device; and in actual production, raw materials must be switched frequently and process parameters must be adjusted frequently, which greatly increases the difficulty of device operation and produces a large number of unqualified products.

[0007] Shell has announced a process for producing cyclopentane rubber filler oil by using high-pressure steam stripping to directly remove H2S and NH3 from the hydrotreated product, which then enters the precious metal hydrodewaxing and supplementary refining reactors. It has three industrial applications in China, but since the cold high-fraction oil containing H2S and NH3 without deep dearomatization is directly incorporated into the distillation system, all products have poor color and stability. Summary of the invention

[0008] In view of the shortcomings of the prior art, the present invention provides a method for producing a high viscosity index lubricating oil base oil by hydrogenation. The present invention uses one or more waxy oils selected from paraffin-based vacuum distillate oil, light deoiled oil, and shale oil vacuum distillate oil as raw materials, and produces a high viscosity index lubricating oil base oil by a two-stage process of hydrogenation treatment-high pressure hydrogen stripping / isomerization dewaxing-supplementary refining.

[0009] A method for producing a high viscosity index lubricating oil base oil by hydrogenation of the present invention comprises the following contents:

[0010] (1) One or more waxy oils selected from paraffin-based vacuum distillate, light deoiled oil, and shale oil vacuum distillate are mixed with hydrogen and then enter a hydroprocessing reactor, and are sequentially contacted with a hydrogenation protection catalyst, a hydrorefining catalyst, and a hydro-modification catalyst for hydrogenation reaction, wherein the hydro-modification catalyst contains amorphous silica-alumina and modified Y molecular sieve;

[0011] (2) The hydrotreatment effluent obtained in step (1) enters a high-pressure stripping tower from the top, hydrogen is injected into the bottom of the tower as a stripping medium, and bottom oil from which H2S and NH3 are removed is obtained at the bottom of the tower, and effluent gas containing H2S, NH3, hydrogen and part of light oil is obtained at the top of the tower;

[0012] (3) The effluent obtained from the top of step (2) containing H2S, NH3, hydrogen and part of the light oil enters a cold high-pressure separator for gas-liquid separation after heat exchange. The cold high-fraction gas phase is selectively desulfurized and then re-enters the hydrotreating reactor as circulating hydrogen. The cold high-fraction oil containing H2S and NH3 enters a cold low-fraction to separate the low-fraction gas. The cold low-fraction oil enters a H2S removal stripping tower to remove H2S and NH3. After pressure increase and heat exchange, it is divided into two streams, one entering the isomerization dewaxing reactor and the other entering the supplementary refining reactor.

[0013] (4) the high-pressure stripping tower bottom oil obtained in step (2) is mixed with hydrogen and then sequentially enters the isomerization dewaxing reactor and the supplementary refining reactor together with the cold low-fraction oil from which H2S and NH3 are removed in step (3) to perform isomerization dewaxing reaction and deep dearomatization reaction to reduce the product freezing point and improve stability;

[0014] (5) The reaction effluent obtained in step (4) is subjected to gas-liquid separation, and the liquid enters the fractionation system to obtain a series of high viscosity index lubricating oil base oils and at least one light white oil with pour point meeting the requirements, small viscosity index loss and good stability.

[0015] Furthermore, the high-pressure stripping tower of step (2) is filled with fillers, such as rolled hole θ ring fillers, and the selection of the fillers is common knowledge of those skilled in the art. Preferably, a ZnO bed is also provided at the lower part of the packing area, and the ZnO bed can be used to remove a small amount of H2S that is not stripped cleanly due to operation fluctuations. The operating conditions of the high-pressure stripping tower are well known to those skilled in the art. For example, the stripping conditions include: a stripping temperature of 150 to 400°C, preferably the outlet temperature of the hydroprocessing reactor, a stripping pressure that is the same as the pressure of the hydroprocessing reactor, and a hydrogen consumption that is higher than the chemical hydrogen consumption of the hydroprocessing, preferably equivalent to the sum of the chemical hydrogen consumption of the hydroprocessing and the dissolved hydrogen of the hydroprocessing system.

[0016] Furthermore, the waxy oil may be one or more of paraffin-based vacuum distillate oil, light deoiled oil, and shale oil vacuum distillate oil. The density of the waxy oil raw material at 20°C is generally 0.900-0.935 g / cm 3 The viscosity index of the waxy oil is generally 85 to 160, and the preferred viscosity index is 115 to 145. The distillation range of the waxy oil raw material is any range consisting of any two values ​​within the range of 140 to 600°C. Preferably, the initial distillation point of the waxy oil raw material is 140 to 350°C, and the final distillation point is 400 to 600°C.

[0017] Furthermore, the hydrorefining catalyst described in step (1) is a conventional hydrorefining catalyst in the art, such as a diesel hydrorefining catalyst or a hydrocracking pretreatment catalyst. The hydrorefining catalyst generally contains Group VIB and / or Group VIII metals as active components, with alumina or silicon-containing alumina as a carrier, the Group VIB metals are generally Mo and / or W, and the Group VIII metals are generally Co and / or Ni. Based on the weight of the catalyst, the content of the Group VIB metal as oxide is 10wt% to 50wt%, and the content of the Group VIII metal as oxide is 3 to 15wt%; its properties are as follows: specific surface area is 100 to 350m 2 / g, pore volume is 0.15~0.6mL / g. Commercial catalysts that can be selected include 3936, 3996, FF-16, FF-26, FF-36, FF-46, FF-56, FF-66 and other hydrofining catalysts developed by Fushun Petrochemical Research Institute of Sinopec. It can also be a catalyst with similar functions developed by foreign catalyst companies, such as UOP's HC-K, HC-P catalysts, Topsoe's TK-555, TK-565 catalysts, and AKZO's KF-847, KF-848, etc.

[0018] Furthermore, the hydro-modification catalyst contains amorphous silica-alumina and modified Y molecular sieve. Based on the weight of the catalyst, the hydro-modification catalyst contains 20wt% to 60wt% amorphous silica-alumina, 5wt% to 25wt% modified Y zeolite, 10wt% to 30wt% of Group VIB metal (in terms of oxide) and 4wt% to 15wt% of Group VIII metal (in terms of oxide). The specific surface area of ​​the hydro-modification catalyst is generally 220 to 300m 2 / g, pore volume is 0.3-0.6mL / g, pore volume with pore diameter of 3-10nm accounts for 75%-95% of the total pore volume, preferably 85%-95%, infrared acidity is 0.30-0.5mmol / g. The properties of the amorphous silica-alumina used are as follows: containing 10wt%-60wt% of silicon oxide, specific surface area of ​​400-650m 2 / g, pore volume of 1.0-1.8mL / g, infrared acidity of 0.34-0.50mmol / g, pore volume of pores with a pore diameter of 4-10nm accounts for 85%-95% of the total pore volume, and pore volume of >15nm accounts for less than 5% of the total pore volume; the preferred properties are as follows: containing 10wt%-35wt% of silicon oxide, specific surface area of ​​530-650m 2 / g, pore volume is 1.2-1.5mL / g. The modified Y zeolite has the following properties: SiO2 / Al2O3 molar ratio is 40-60, unit cell constant is 2.425-2.440nm, relative crystallinity is 80-100%, infrared acidity is 0.1-0.5mmol / g, among which medium-strong acid at 250-550℃ is concentrated, accounting for 60-70% of the total acid, and specific surface area is 600-900m 2 / g, pore volume 0.3~0.6mL / g, of which the pore volume of secondary mesopores of 4~15nm accounts for 40%~50% of the total pore volume. The hydrogenation reforming catalyst may also contain one or more components of aluminum oxide, zirconium oxide and titanium oxide.

[0019] Furthermore, in step (3), the weight ratio of the liquid obtained from the H2S removal stripping tower entering the isodewaxing reactor and the supplementary refining reactor is generally 1:9 to 9:1, preferably 1:4 to 4:1.

[0020] Further, the isodewaxing reactor is loaded with an isodewaxing catalyst. The isodewaxing catalyst used in the method of the present invention can be selected from the commonly used lubricating oil hydroisomerization catalyst in the art, and a commercial hydroisomerization catalyst can be used, or it can be prepared according to general knowledge in the art. The hydroisomerization catalyst carrier is generally alumina and TON structure molecular sieve. TON structure molecular sieve is a group of substances composed of NU-10 molecular sieve, ZSM-22 molecular sieve, ZSM-23 molecular sieve, ZSM-48 molecular sieve, etc., preferably ZSM-22 molecular sieve. The content of TON structure molecular sieve in the catalyst is generally 30wt%~80wt%, preferably 40wt%~70wt%, and part of silicon oxide can also be added to the carrier; the active metal component is one or more of Pt, Pd, Ru and Rh, and the content in the catalyst is 0.1wt%~5.0wt%; the optional auxiliary component is one or more of boron, fluorine, chlorine and phosphorus, and the content in the catalyst is 0.1wt%~5.0wt%. The specific surface area of ​​the hydroisomerization catalyst is 150~500m 2 / g, and the pore volume is 0.15~0.60ml / g.

[0021] Furthermore, a supplementary refining catalyst is loaded in the supplementary refining reactor. The supplementary refining catalyst is a conventional reduced hydrogenation refining catalyst, and its active metal is one or two of Pt and Pd, or reduced nickel. The weight content of the active metal in the precious metal catalyst is generally 0.05% to 1% in terms of metal, and the weight content of nickel in the reduced nickel catalyst is 30% to 80% in terms of oxide. The carrier of the supplementary refining catalyst is generally Al2O3 or Al2O3-SiO2, and may also contain at least one of P, Ti, B, Zr, etc. as an auxiliary agent. The supplementary refining catalyst needs to be conventionally reduced before use to ensure that the hydrogenation active metal is in a reduced state during the reaction. The supplementary refining catalyst can be selected from ordinary commercial catalysts in this field, or prepared according to ordinary methods in this field.

[0022] Further, the operating conditions of the hydroprocessing reactor are: hydrogen partial pressure 5.0-18.0 MPa, preferably 10.0-16.0 MPa; reaction temperature 230-430°C, preferably 350-410°C; hydrogen-to-oil volume ratio 200-1500, preferably 600:1-1000:1; volume space velocity 0.1-10.0 h -1 , preferably 0.3 h -1 ~0.5h -1 In the hydroprocessing reaction zone, the loading ratio of the hydrofining catalyst and the hydromodification catalyst is 1:9 to 9:1, preferably 4:6 to 6:4. The hydroprocessing reaction zone may further include a hydroprotection catalyst. The hydroprotection catalyst is usually arranged upstream of the hydrofining catalyst. The amount of the hydroprotection catalyst is well known to those skilled in the art.

[0023] Further, the reaction conditions of the hydroisomerization dewaxing reactor are: reaction temperature of 300°C to 380°C, preferably 310°C to 340°C, hydrogen partial pressure of 2.0MPa to 18.0MPa, preferably 10.0 to 15.0MPa, and raw oil liquid volume space velocity of 0.2h -1 ~6.0h -1 , preferably 0.4 h -1 ~1.0h -1 , the volume ratio of hydrogen to oil is 400:1 to 1500:1, preferably 600:1 to 800:1.

[0024] Further, the reaction conditions of the supplementary refining reactor are: reaction temperature is 200°C to 300°C, preferably 210°C to 280°C, hydrogen partial pressure is 6.0MPa to 18.0MPa, preferably 10.0 to 15.0MPa, volume space velocity is 0.2h -1 ~3.0h -1 , preferably 0.3h -1 ~0.6h -1 , the volume ratio of hydrogen to oil is 400:1 to 1500:1, preferably 600:1 to 800:1.

[0025] Compared with the prior art, the method of the present invention has the following advantages:

[0026] 1. The present invention adopts an improved two-stage hydrogenation process of high-pressure hydrogen stripping, hydrogenation treatment, high-pressure hydrogen stripping / isomerization decondensation, and deep supplementary refining, which solves the problem that two sets of fractionation systems need to be set up in the traditional two-stage hydrogenation process to eliminate the influence of H2S and NH3 on the precious metal catalyst. Most of the reaction feed (more than 85w%) of the improved two-stage hydrogenation process of the present invention is a one-time pressure-increasing process, which avoids the process of pressure-increasing, pressure-reducing, and re-pressurizing of all raw oil between the two reaction zones of the traditional two-stage process device, saving electricity consumption; at the same time, the stripping hydrogen injected at the bottom of the high-pressure stripping tower also enters the circulating hydrogen system of the hydroprocessing device as part of the supplementary hydrogen. By adjusting the operating temperature of the high-pressure hydrogen stripping tower, the bottom effluent of the stripping tower can directly enter the isomerization dewaxing reactor, and there is no need to set up an isomerization dewaxing feed heating furnace, which also avoids the energy loss in the process of heating, cooling, and reheating of all reaction products in the traditional two-stage process, saving electricity consumption and fuel gas consumption, and saving land.

[0027] 2. The effluent from the top of the high-pressure hydrogen stripping tower enters the cold high-pressure separator after heat exchange. After separating the gas phase circulating hydrogen, the cold high-fraction oil containing H2S and NH3 is depressurized and enters the cold low-pressure separator to separate the low-fraction gas. The cold low-fraction oil enters the conventional H2S removal stripping tower to remove H2S and NH3. After heat exchange and temperature increase, it is divided into two streams: one stream enters the isomerization dewaxing reactor, and the other stream enters the supplementary refining reactor. This solves the problem of the Shell process cold low-fraction oil directly entering the fractionation system to affect the product color and stability; at the same time, by adjusting the ratio of the two streams, it also solves the problem that when the pour point of the heavy lubricating oil base oil is qualified, the pour point of the light lubricating oil base oil is greatly excessive and the viscosity index loss is large, thereby improving the liquid recovery and the yield of the target product. BRIEF DESCRIPTION OF THE DRAWINGS

[0028] Figure 1 It is a schematic diagram of the process flow of the method of the present invention. DETAILED DESCRIPTION

[0029] The method of the present invention is described in more detail below with reference to the accompanying drawings and specific embodiments.

[0030] Combination Figure 1 The method for producing high viscosity index lubricating oil base oil by hydrogenation of the present invention is as follows: one or more waxy oils selected from paraffin-based vacuum distillate oil, light deoiling oil, and shale oil vacuum distillate oil are mixed with circulating hydrogen from pipeline 2 through pipeline 1, and then first enter a hydroprocessing reactor 3 to undergo desulfurization, denitrogenation, deoxygenation, and aromatic saturation reactions; subsequently, a hydroprocessing product 4 enters a high-pressure hydrogen stripping tower 5, and stripping hydrogen 6 is blown into the bottom of the tower to obtain a tower bottom oil 7 from which H2S and NH3 are removed; the top of the high-pressure stripping tower enters a separation system 8 of a hydroprocessing section, and the separation system 8 generally includes a high-pressure separator and a low-pressure separator; the separated gas phase is returned to the hydroprocessing reaction system as circulating hydrogen through pipeline 2, and the liquid phase product enters a H2S removal stripping tower 10 through pipeline 9 to remove H2S and NH3. After the tower bottom oil 11 is pressurized and heat exchanged, it is divided into two streams 13 and 14 according to a certain ratio, and enters the isomerization dewaxing reactor 15 and the supplementary refining reactor 17 respectively. The high-pressure stripping tower bottom oil 7 is mixed with the stream 13 and the new hydrogen 12 and directly enters the isomerization dewaxing reactor 15 to carry out the isomerization decondensation reaction. The isomerization dewaxing product 16 is mixed with the stream 14 and enters the supplementary refining reactor 17 to carry out the deep dearomatization reaction to improve the light and thermal stability of the product. The supplementary refining reaction product 18 enters the separation system 19, which generally includes a high-pressure separator and a low-pressure separator. The low-pressure oil enters the fractionation system 20, which includes a normal pressure fractionation tower and a vacuum fractionation tower to obtain light white oil 21, industrial white oil 22 and a series of high viscosity index lubricant base oil products 23.

[0031] The specific contents and effects of the method of the present invention are described below by means of specific embodiments.

[0032] The following examples will further illustrate the method provided by the present invention, but the scope of the present invention is not limited thereby. The properties of the raw materials treated by the present invention are shown in Table 1, and the physical and chemical properties of the catalyst used are shown in Table 2.

[0033] Table 1 Raw material properties

[0034]

[0035] Table 2 Catalyst

[0036]

[0037] Example 1

[0038] The raw material is paraffin-based three-line oil 1 listed in Table 1. The physical and chemical properties of the catalyst are shown in Table 2. The hydroprocessing zone is filled with a series of hydrogenation protective agents, hydrorefining catalysts and hydroreforming catalysts, wherein the series of hydrogenation protective agents: hydrorefining catalysts: hydroreforming catalysts = 1:5.5:3.5 (volume). The principle flow chart is shown in Figure 1 The results are shown in Tables 3~4.

[0039] Example 2

[0040] The raw material is paraffin-based three-line oil 2 listed in Table 1, and the same catalyst as in Example 1 is used. The principle flow chart is shown in Figure 1 The results are shown in Tables 3~4.

[0041] Comparative Example 1

[0042] Comparative Example 1 adopts the traditional two-stage hydrogenation process. The process flow is as follows: the hydroprocessed product first enters the hot high-pressure separator, and then enters the cold high-pressure separator after gas phase heat exchange. The cold high-fraction gas phase enters the circulating hydrogen system, and the cold high-fraction oil enters the cold low-fraction after decompression, and the top of the separator is the low-fraction gas; the hot high-fraction oil enters the hot low-fraction after decompression, and the top is the low-fraction gas; the hot low-fraction oil (most of the product) and the cold low-fraction oil are mixed and then all enter the dehydrogen sulfide stripping tower (pressure is about 1MPa), and a small amount of hydrogen sulfide gasoline is discharged from the top of the tower and all liquids (sulfur content <10µg.g -1 , nitrogen content <2µg.g -1 ) After heating and pressure increase, it enters the second stage isomerization dewaxing-supplementary refining stage; the supplementary refined product is subjected to gas-liquid separation and fractionation to obtain various products.

[0043] Comparative Example 1 uses the same raw materials and the same catalyst as in Example 1. The process conditions and test results are shown in Table 3-4.

[0044] Comparative Example 2

[0045] Comparative Example 2 adopts the Shell process ("Research on the Application of High-Pressure Hydrogen Stripping in the Production Process of Fully Hydrogenated Lubricant Oil", Xue Nan, "Petroleum Refining and Chemical Industry", October 2019).

[0046] Comparative Example 2 uses the same raw materials and catalyst as in Example 1. The process is that the hydrotreated product directly enters the isomerization dewaxing-supplementary refining section after high-pressure stripping, the hot high-fraction gas phase enters the cold low-pressure separator, the cold low-fraction gas is returned to the reaction system as circulating hydrogen, and the cold low-fraction oil directly enters the subsequent fractionation system for fractionation to obtain various products. The results are shown in Tables 3-4.

[0047] Table 3 Process conditions

[0048]

[0049]

[0050] Table 4 Test results

[0051]

[0052] From the data in Table 3-4, it can be seen that by using paraffin-based vacuum distillate as raw material and adopting the two-stage process of hydroprocessing-high-pressure hydrogen stripping / isodewaxing-supplementary refining, products of equivalent quality to the conventional two-stage process of hydroprocessing / isodewaxing-supplementary refining can be produced. At the same time, the energy consumption of the conventional two-stage process of all feed pressure-reducing-pressure-increasing, heating-cooling-heating processes is avoided, the equipment area is reduced, and the energy consumption of the equipment is reduced by about 12%. In addition, since the cold high-fraction oil is subjected to conventional stripping and enters the isodewaxing reactor and the supplementary refining reactor respectively according to the optimized ratio, the viscosity index loss of the light lubricating oil base oil is reduced while ensuring that the pour point of the series of lubricating oil base oils is qualified, the light oil loss at the top of the conventional two-stage distillation process H2S stripping tower is reduced, and the liquid recovery and target product yield are improved. Although the Shell process has lower energy consumption, the oxidation stability of the entire line of products is unqualified because the cold low-fraction oil is directly fed into the fractionation system without hydrogenation supplementary refining.

Claims

1. A method for producing a high viscosity index lubricating oil base oil by hydrogenation, comprising the following: (1) One or more waxy oils selected from paraffin-based vacuum distillate, light deoiled oil, and shale oil vacuum distillate are mixed with hydrogen and then enter a hydroprocessing reactor, and are sequentially contacted with a hydrogenation protection catalyst, a hydrorefining catalyst, and a hydro-modification catalyst for hydrogenation reaction, wherein the hydro-modification catalyst contains amorphous silica-alumina and modified Y molecular sieve; (2) The hydrotreatment effluent obtained in step (1) enters a high-pressure stripping tower from the top, hydrogen is injected into the bottom of the tower as a stripping medium, and bottom oil from which H2S and NH3 are removed is obtained at the bottom of the tower, and an effluent containing H2S, NH3, hydrogen and part of light oil is obtained at the top of the tower; (3) The effluent obtained from the top of step (2) containing H2S, NH3, hydrogen and part of the light oil enters a cold high-pressure separator for gas-liquid separation after heat exchange, and the cold high-fraction oil containing H2S and NH3 enters a cold low-fraction separator to separate the low-fraction gas. The cold low-fraction oil enters a H2S removal stripping tower to remove H2S and NH3, and is divided into two streams after pressure increase and heat exchange, one of which enters an isomerization dewaxing reactor and the other enters a supplementary refining reactor; (4) the high-pressure stripping tower bottom oil obtained in step (2) is mixed with hydrogen and then enters the isomerization dewaxing reactor and the supplementary refining reactor in sequence together with the cold low-fraction oil from which H2S and NH3 are removed in step (3) to perform isomerization dewaxing reaction and deep dearomatization reaction; (5) The reaction effluent obtained in step (4) is subjected to gas-liquid separation, and the liquid enters a fractionation system to obtain a high viscosity index lubricating oil base oil and at least one light white oil; in, The stripping pressure of the high-pressure stripping tower in step (2) is the same as the pressure of the hydroprocessing reactor. The reaction pressure of the hydroprocessing reactor is 5.0 to 18.0 MPa in terms of hydrogen partial pressure.

2. The method according to claim 1, characterized in that The high-pressure stripping tower is filled with fillers, and the fillers are selected from rolled hole θ ring fillers.

3. The method according to claim 2, characterized in that A ZnO bed is arranged in the lower part of the packing zone.

4. The method according to claim 1, characterized in that The stripping temperature of the high-pressure stripping tower is 150-400° C., and the amount of hydrogen used is higher than the chemical hydrogen consumption of the hydrogenation treatment.

5. The method according to claim 4, characterized in that The stripping temperature of the high-pressure stripping tower is the outlet temperature of the hydroprocessing reactor, and the amount of hydrogen used is equivalent to the sum of the chemical hydrogen consumption of the hydroprocessing and the dissolved hydrogen of the hydroprocessing system.

6. The method according to claim 1, characterized in that The cold high-fraction gas phase obtained in step (3) is selectively desulfurized and then re-enters the hydroprocessing reactor as circulating hydrogen.

7. The method according to claim 1, characterized in that The density of the wax-containing oil at 20°C is 0.900-0.935 g / cm 3 .

8. The method according to claim 1, characterized in that The viscosity index of the waxy oil is 85-160.

9. The method according to claim 1, characterized in that The initial distillation point of the wax-containing oil is 140-350°C, and the final distillation point is 400-600°C.

10. The method according to claim 1, characterized in that Based on the weight of the catalyst, the hydrogenation and reforming catalyst contains 20wt% to 60wt% of amorphous silicon aluminum, 5wt% to 25wt% of modified Y molecular sieve, 10wt% to 30wt% of Group VIB metal in terms of oxide, and 4wt% to 15wt% of Group VIII metal in terms of oxide.

11. The method according to claim 10, characterized in that The specific surface area of ​​the hydrogenation catalyst is 220 to 300 m 2 / g, the pore volume is 0.3-0.6mL / g, the pore volume with a pore diameter of 3-10nm accounts for 75%-95% of the total pore volume, and the infrared acidity is 0.30-0.5mmol / g.

12. The method according to claim 1, characterized in that In step (3), the weight ratio of the liquid obtained from the H2S stripping tower entering the isomerization dewaxing reactor and the supplementary refining reactor is 1:9~9:

1.

13. The method according to claim 1, characterized in that The isodewaxing reactor is filled with an isodewaxing catalyst, the carrier of the isodewaxing catalyst is alumina and TON structure molecular sieve, the active metal component is one or more of Pt, Pd, Ru and Rh, and the optional auxiliary component is one or more of boron, fluorine, chlorine and phosphorus.

14. The method according to claim 1, characterized in that The supplementary refining reactor is filled with a supplementary refining catalyst, wherein the supplementary refining catalyst is a reduced hydrorefining catalyst, and the active metal is one or both of Pt and Pd, or reduced nickel.

15. The method according to claim 1, characterized in that The operating conditions of the hydroprocessing reactor are: reaction temperature 230-430°C, hydrogen-to-oil volume ratio 200-1500, volume space velocity 0.1-10.0h -1 .

16. The method according to claim 1, characterized in that The reaction conditions of the isomerization dewaxing reactor are: reaction temperature of 300°C to 380°C, hydrogen partial pressure of 2.0MPa to 18.0MPa, raw oil liquid volume space velocity of 0.2h -1 ~6.0h -1 , hydrogen-oil volume ratio 400:1~1500:

1.

17. The method according to claim 1, characterized in that The reaction conditions of the supplementary refining reactor are: reaction temperature of 200°C to 300°C, hydrogen partial pressure of 6.0MPa to 18.0MPa, volume space velocity of 0.2h -1 ~3.0h -1 , hydrogen-oil volume ratio 400:1~1500:1.

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