A method for producing industrial white oil by hydrogenation

Abstract

The invention discloses a method for producing industrial white oil by hydrogenation. Paraffin-based, intermediate-based or cycloalkyl straight-run diesel is used as raw material, and firstly, hydrofining is performed; the hydrofining effluent is stripped by medium-pressure hydrogen to remove impurities such as H2S and NH3 that affect the reduction catalyst, and then enters the supplementary refining-isomerization dewaxing-post-refining reaction zone to reduce the pour point and aromatic content of the oil product and improve the color. The method of the invention can obtain high-quality Class II industrial white oil, while producing a small amount of light white oil as a by-product, and greatly reduces the energy consumption caused by the two-stage method of boosting-depressurization-boosting, and heating-cooling-heating, and has the characteristics of simple process and small footprint.

Description

Technical Field

[0001] The invention belongs to the technical field of hydrogenation and relates to a method for producing industrial white oil by hydrogenation. Specifically, the invention uses paraffin-based, intermediate-based or cycloalkyl straight-run diesel as raw materials and adopts a medium-pressure two-stage hydrogenation process of hydrogenation refining-medium-pressure hydrogen stripping-precious metal supplementary refining-isomerization dewaxing to produce Class II industrial white oil that meets the NB / SH / T 0006-2017 standard. Background Art

[0002] White oil is a colorless, tasteless, odorless and non-corrosive special mineral oil that has been ultra-deeply refined (removing impurities such as sulfur, nitrogen and aromatics). White oil is basically composed of saturated hydrocarbons and belongs to the lubricating oil fraction. It has good chemical inertness and excellent light and heat stability. White oil is widely used in daily chemical industry, pharmaceutical production, food processing, fiber and textile, polystyrene resin, petrochemical industry, plastic and rubber processing, leather processing, instrumentation and electricity, agriculture and other fields.

[0003] There are two main production processes for white oil: sulfonation and hydrogenation. At the end of the 19th century, foreign countries began to produce white oil using concentrated sulfuric acid or oleum sulfonation. The white oil produced by this method was unstable in quality, low in yield, large in raw material loss, large in acid consumption, and difficult to handle in large amounts of acid residue, causing serious environmental pollution. Later, the sulfur trioxide sulfonation process emerged, which greatly reduced acid consumption and increased white oil yield, but the environmental pollution problem was still not solved.

[0004] Since the 1960s, hydrogenation refining technology has been applied to the production process of white oil. Since hydrogenation can remove trace amounts of sulfur, nitrogen and oxygen in oil, and convert aromatics into cycloalkanes, making them white oil components, the yield of white oil products is high, thus improving the yield and eliminating the generation of three wastes. At present, white oil produced by hydrogenation at home and abroad accounts for more than 90% of the total white oil.

[0005] Depending on the different raw materials and product requirements, one-stage hydrogenation or multi-stage hydrogenation is used. Hydrocracked diesel or tail oil with low sulfur and nitrogen content, generally with an aromatic content of less than 5%, can be hydrogenated to obtain industrial grade white oil or even food and pharmaceutical grade white oil. Straight-run diesel with high sulfur and nitrogen content, with an aromatic content of up to 10% to 20%, generally requires two-stage hydrogenation to obtain Class II industrial white oil that meets the NB / SH / T 0006-2017 standard.

[0006] In recent years, diesel sales have been difficult. Considering the material balance of the whole plant, it is very difficult to simply reduce the diesel production. Using the existing medium-pressure diesel hydrogenation unit and using one or two-stage precious metal catalysts to produce industrial white oil is a good choice, but it is also affected by the existing layout of the unit and difficult to transform. At the same time, for most straight-run diesel, the production of industrial white oil through the existing process still has the problem of product pour point: for the No. 5 industrial white oil in the range of 270~350℃, the pour point is mostly above 0℃, and there is also a problem of pour point reduction.

[0007] CN201310540349.X discloses a method for producing special lubricating base oil. After the wax oil raw material is mixed with hydrogen, it passes through at least two serially connected hydrogenation reaction zones in sequence. According to the material flow direction, each hydrogenation reaction zone includes a hydrodecondensation catalyst bed and a refined decondensation catalyst composite bed filled with a hydrodecondensation catalyst and a hydrorefining catalyst. The reaction effluent obtained from the last hydrogenation reaction zone is separated and fractionated to obtain lubricating base oil, white oil and other products. The inventive method reasonably combines and utilizes the temperature drop in the hydrodecondensation process and the temperature rise in the hydrogenation process. While producing low-condensation-point lubricating base oil, it reduces the hot spot temperature of the device and prolongs the operation cycle. However, there are also problems such as high operating pressure, high energy consumption, and inability to use existing devices for transformation. Summary of the invention

[0008] In view of the shortcomings of the prior art, the present invention provides a method for producing industrial white oil by hydrogenation. Specifically, a method for producing Class II industrial white oil that meets the NB / SH / T 0006-2017 standard by using paraffin-based, intermediate-based or cycloalkyl straight-run diesel as raw materials and adopting a medium-pressure two-stage hydrogenation process of hydrofining-medium-pressure hydrogen stripping-precious metal supplementary refining-isomerization dewaxing.

[0009] A method for producing industrial white oil by hydrogenation of the present invention comprises the following contents:

[0010] (1) Paraffin-based, intermediate-based or cycloalkyl straight-run diesel is mixed with hydrogen and enters the hydrotreating reaction zone, where it contacts the hydrotreating catalyst in turn to undergo a hydrotreating reaction to remove impurities such as sulfur, nitrogen and oxygen from the raw material and saturate some aromatics;

[0011] (2) The hydrotreating effluent obtained in step (1) enters a medium-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 gas effluent obtained in 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 de-H2Sed and 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 de-H2S stripping tower to remove H2S and NH3, and enters a supplementary refining reactor after pressure increase and heat exchange.

[0013] (4) The tower bottom oil obtained in step (2) from which H2S and NH3 have been removed is mixed with hydrogen and then enters the supplementary refining reactor together with the cold low fraction oil from which H2S and NH3 have been removed, and first contacts and reacts with the upper reduced supplementary refining catalyst to deeply saturate aromatic hydrocarbons, and then reacts with the lower isomerization dewaxing (depressing pour point) catalyst at a higher temperature to reduce the pour point of the product by virtue of the temperature rise of the dearomatization reaction, and then reacts with the lower reduced supplementary refining catalyst to saturate a small amount of olefins produced during the isomerization dewaxing process;

[0014] (5) The reaction effluent obtained in step (4) enters a fractionation system to obtain Class II industrial white oil that meets the NB / SH / T 0006-2017 standard and a small amount of a series of light white oils that meet the technical index requirements of light white oil (II) in NB / SH / T 0913-2015.

[0015] Furthermore, the straight-run diesel can be at least one of a paraffin-based, intermediate-based or naphthenic straight-run diesel fraction and a catalytic cracking diesel fraction. The aromatic content of the straight-run diesel is generally 1 to 99 wt%, preferably 20 to 35 wt%, and the density at 20°C is generally 0.825 to 0.915 g / cm 3 The distillation range of straight-run diesel can be any range formed by any two point values ​​within the range of 140-400°C; preferably, the initial distillation point of straight-run diesel is 160-200°C, and the final distillation point is 300-400°C.

[0016] Furthermore, the hydrotreating reaction zone also includes a hydrotreating protection catalyst, which is usually arranged upstream of the hydrotreating catalyst.

[0017] Furthermore, in step (2), the medium-pressure stripping tower is filled with fillers, such as rolled hole θ ring fillers, and the selection of the fillers is common knowledge for those skilled in the art. Preferably, a ZnO bed is also provided at the lower part of the filler zone, 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 medium-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 hydrofining reactor, a stripping pressure that is the same as the pressure of the hydrofining reaction zone, and a hydrogen consumption that should be higher than the chemical hydrogen consumption of the hydrofining reaction zone; preferably, the hydrogen consumption is equivalent to the sum of the chemical hydrogen consumption of the hydrofining and the dissolved hydrogen of the hydrofining system.

[0018] 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 has Group VIB and / or Group VIII metals as active components and alumina or silicon-containing alumina as a carrier. The Group VIB metal is generally Mo and / or W, and the Group VIII metal is generally Co and / or Ni. Based on the weight of the catalyst, the content of the Group VIB metal in oxide form is 10wt% to 50wt%, and the content of the Group VIII metal in oxide form is 3 to 15wt%; its physical properties are as follows: specific surface area is 100 to 350m 2 / g, pore volume is 0.15~0.6mL / g. Commercial hydrotreating catalysts that can be selected include 3FF-26, FF-36, FF-46, FF-56, FF-66, FHUDS-2, FHUDS-5 and other hydrotreating 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 HC-K and HC-P catalysts of UOP, TK-555 and TK-565 catalysts of Topsoe, and KF-847 and KF-848 of AKZO.

[0019] Furthermore, the process conditions of the hydrofining in step (1) include: hydrogen partial pressure of 5.0 to 18.0 MPa, preferably 5.0 to 10.0 MPa; reaction temperature of 230 to 430°C, preferably 300 to 410°C; hydrogen to oil volume ratio of 200 to 1500, preferably 500:1 to 1000:1; volume space velocity of 0.1 to 10.0 h -1 , preferably 0.3 h -1 ~1.0h -1 .

[0020] Furthermore, the isomerization dewaxing catalyst can select the commonly used lubricating oil hydroisomerization catalyst in the art, and can use commercial hydroisomerization catalysts or can be prepared according to general knowledge in the art. The hydroisomerization catalyst includes a carrier and an active metal supported on the carrier. The hydroisomerization catalyst carrier generally includes alumina and TON structure molecular sieve. The TON structure molecular sieve is selected from at least one of NU-10 molecular sieve, ZSM-22, 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 30wt%~80wt%, preferably 30wt%~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 dewaxing catalyst is 150~500m 2 / g, and the pore volume is 0.15~0.60ml / g.

[0021] Furthermore, in the supplementary refining reactor, the loading volume ratio of the supplementary refining catalyst, the isodewaxing catalyst and the post-refining catalyst is well known to those skilled in the art, and the volume ratio is generally 1~5:1~5:1~5.

[0022] Furthermore, the supplementary refining catalyst and the post-refining catalyst are conventional reduced hydrogenation refining catalysts, and the active metals thereof are one or both of Pt and Pd, or reduced nickel. The weight content of active metals in precious metal catalysts is generally 0.05% to 1%; in reduced nickel catalysts, the weight content of nickel in terms of oxide is 30% to 80%. The catalyst carrier is generally Al2O3 or Al2O3-SiO2, and the catalyst may also contain at least one of P, Ti, B, Zr, etc. as an auxiliary agent. Before use, the supplementary refining (post-refining) catalyst needs to be conventionally reduced to ensure that the hydrogenation active metal is in a reduced state during the reaction. The supplementary refining catalyst or the post-refining catalyst can be selected from ordinary commercial catalysts in this field, or prepared according to ordinary methods in this field.

[0023] Further, the reaction conditions of the supplementary refining reaction zone are as follows: reaction temperature is 100°C to 300°C, preferably 150°C to 250°C, hydrogen partial pressure is 3.0MPa to 18.0MPa, preferably 5.0 to 8.0MPa, volume space velocity is 0.2h -1 ~3.0h -1 , preferably 0.3h -1 ~0.8h -1 , the volume ratio of hydrogen to oil is 400:1 to 1500:1, preferably 400:1 to 600:1.

[0024] Further, the reaction conditions of the isomerization dewaxing section in the supplementary refining reaction zone are: reaction temperature is 100°C to 350°C, preferably 250°C to 320°C, hydrogen partial pressure is 3.0MPa to 18.0MPa, preferably 5.0 to 8.0MPa, volume space velocity is 0.2h -1 ~5.0h -1 , preferably 2.0h -1 ~3.0h -1 , the volume ratio of hydrogen to oil is 400:1 to 1500:1, preferably 400:1 to 600:1.

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

[0026] 1. The present invention adopts a medium-pressure two-stage hydrogenation process of hydrofining-medium-pressure hydrogen stripping-precious metal supplementary refining / isomerization dewaxing to produce industrial white oil, avoiding the design scheme of setting two sets of fractionation systems in the traditional two-stage hydrogenation process to avoid the influence of H2S and NH3 on the precious metal catalyst. At the same time, in the medium-pressure two-stage hydrogenation process of the present invention, most of the reaction feed (~80w%) is a one-time pressure-increasing process, avoiding 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 power consumption. At the same time, the stripping hydrogen injected at the bottom of the medium-pressure stripping tower enters the circulating hydrogen system of the hydrofining device as part of the supplementary hydrogen; by adjusting the operating temperature of the medium-pressure hydrogen stripping tower, the bottom effluent of the stripping tower can directly enter the supplementary refining reactor without setting a supplementary refining feed heating furnace, avoiding the energy loss in the process of heating, cooling, and re-heating of all reaction products in the traditional two-stage process, saving power consumption and fuel gas consumption, and also saving land.

[0027] 2. In the present invention, the top effluent of the medium-pressure stripping tower enters a cold high-pressure separator after heat exchange, and after separating the gas phase circulating hydrogen, the cold high-fraction oil containing H2S and NH3 enters a cold low-pressure separator after pressure reduction to separate the low-fraction gas, and the cold low-fraction oil enters a conventional H2S removal stripping tower to remove H2S and NH3, and enters a supplementary refining reactor after heat exchange and pressure increase, thereby solving the problem of the cold high-fraction oil directly entering the distillation system and affecting the product color, aromatics content and stability, or the problem of directly exiting the device and having nowhere to go.

[0028] 3. In the present application, the inventor innovatively graded the supplementary refining catalyst and the isodewaxing catalyst: the supplementary refining catalyst is loaded in the upper part of the supplementary refining reactor, and the isodewaxing catalyst and a small amount of supplementary refining catalyst are loaded in the lower part as the post-refining catalyst. The result of this grading method is that the diesel after a period of hydrorefining is first deeply hydrodearomatized in the upper part of the supplementary refining reactor, and then, with the help of the temperature rise of the dearomatization reaction, it contacts and reacts with the downstream isodewaxing catalyst at a higher temperature to reduce the pour point of the product. In the prior art, due to the different reaction temperature ranges of isodewaxing and supplementary refining, there is a large temperature difference, and it is impossible to achieve simultaneous isodewaxing and supplementary refining in one reactor. The present invention sets the isodewaxing catalyst upstream of the supplementary refining catalyst, and the isodewaxing catalyst and the supplementary refining catalyst can be loaded in one reactor, thereby reducing the number of reactors set up. At the same time, the deeply dearomatized diesel is also more conducive to the subsequent isomerization dewaxing reaction, such as reducing the tendency of isomerization catalyst to coke, effectively increasing the operation cycle of isomerization dewaxing catalyst, reducing side reactions, and increasing liquid recovery. At the same time, when the supplementary refining catalyst is selected as Pt-Pd precious metal catalyst, placing it on the upper part of the supplementary refining reactor can also improve the sulfur resistance of the entire supplementary refining system, which is also conducive to increasing the operation cycle of the device. BRIEF DESCRIPTION OF THE DRAWINGS

[0029] Figure 1 It is a process flow diagram of the method of the present invention. DETAILED DESCRIPTION

[0030] The method of the present invention is described in more detail below with reference to the accompanying drawings and specific embodiments.

[0031] Combination Figure 1The method for producing industrial white oil by hydrogenation of the present invention is as follows: straight-run diesel passes through pipeline 1, is mixed with circulating hydrogen from pipeline 2, and enters hydrofining reaction zone 3 to undergo desulfurization, denitrogenation, deoxygenation and aromatic saturation reactions; then the hydrofining product 4 enters a medium-pressure hydrogen stripping tower 5, and stripping hydrogen 6 is blown into the bottom of the tower to obtain bottom oil 7 from which H2S and NH3 are removed; the gas at the top of the medium-pressure stripping tower enters a separation system 8 of the hydrofining 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 hydrofining reaction system as circulating hydrogen through pipeline 2, and the liquid phase product enters a conventional H2S removal stripping tower 1 through pipeline 9. 0, the tower bottom oil 11 from which H2S and NH3 have been removed is pressurized and heat exchanged, mixed with new hydrogen 12 and second-stage circulating hydrogen 13 and enters a supplementary refining reactor 14, the high-pressure stripping tower bottom oil 7 is heat exchanged and mixed with logistics 11 and directly enters the supplementary refining reactor 14, undergoes deep dearomatization, isomerization decondensation and olefin saturation reactions, reduces the aromatic content and improves the color, the supplementary refined product 15 enters the supplementary refining section separation system 16, the separation system 16 generally includes a high-pressure separator and a low-pressure separator, the low-pressure oil enters the fractionation system 17, the fractionation system includes an atmospheric pressure fractionation tower and a vacuum fractionation tower, and obtains light naphtha 18, light white oil 19 that meets the requirements of NB / SH / T 0913-2015, and industrial white oil products 20 that meet the requirements of NB / SH / T 0006-2017.

[0032] The specific contents and effects of the method of the present invention are described below by means of specific embodiments.

[0033] 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.

[0034] Table 1 Raw material properties

[0035]

[0036] Table 2 Physical and chemical properties of catalysts

[0037]

[0038] Example 1

[0039] The raw material is the straight-run diesel 1 listed in Table 1, and the physical and chemical properties of the catalyst are shown in Table 2. The hydrorefining zone is filled with a hydroprotectant and a hydrorefining catalyst, and the filling ratio is a hydroprotectant: hydrorefining catalyst = 1:9 (volume); in the supplementary refining reactor: the volume ratio of the supplementary refining catalyst: the isomerization dewaxing catalyst: the post-refining catalyst is 3:1:2. See the principle flow chart Figure 1 The results are shown in Tables 3~4.

[0040] Example 2

[0041] The raw material is the straight-run diesel 2 listed in Table 1, and the physical and chemical properties of the catalyst are shown in Table 2. The hydrorefining zone is filled with a hydroprotectant and a hydrorefining catalyst, and the filling ratio is a hydroprotectant: hydrorefining catalyst = 1:9 (volume); the volume ratio of the supplementary refining catalyst: isomerization dewaxing catalyst: post-refining catalyst in the supplementary refining reactor is 3:1:2. See the principle flow chart Figure 1 The results are shown in Tables 3~4.

[0042] Comparative Example 1

[0043] The traditional two-stage hydrogenation process is adopted. The process flow is as follows: the hydrorefined 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 of the hot low-fraction 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 conventional desulfurization hydrogen stripping tower (pressure is about 1MPa), a small amount of hydrogen sulfide gasoline is discharged from the top of the tower, and all the liquid at the bottom of the stripping tower (sulfur content <10µg.g -1 , nitrogen content <2µg.g -1 ) After the temperature and pressure are increased, it passes through the second-stage isomerization dewaxing section and the supplementary refining section in sequence; the supplementary refining product is subjected to gas-liquid separation and fractionation to obtain various products.

[0044] Comparative Example 1 uses the same raw materials and the same catalyst as in Example 1. The process conditions and results are shown in Table 3-4.

[0045] Comparative Example 2

[0046] The conventional two-stage process was adopted, and the process flow was the same as that of Comparative Example 1. Comparative Example 2 used the same raw materials and the same catalyst as those in Example 2. The process conditions and test results are shown in Table 3-4.

[0047] Table 3 Process conditions

[0048]

[0049] *, for fresh feed.

[0050] Table 4 Test results

[0051]

[0052] **, for fresh feed.

[0053] From the data in Table 3-4, it can be seen that by using straight-run diesel as raw material and adopting the process of the present invention, a product of equivalent quality to the conventional two-stage hydrofining / supplementary refining process can be produced, while avoiding the energy consumption of all the pressure-raising-reducing-pressurizing, temperature-lowering-temperature-raising processes of the two-stage feed in the conventional two-stage process, reducing the footprint of the device and reducing the energy consumption of the device by about 10%. In addition, since only the cold high-fraction oil enters the supplementary refining reactor through conventional stripping, the light oil loss at the top of the conventional two-stage process H2S stripping tower is greatly reduced, thereby improving the total liquid recovery and the target product yield, and at the same time, it is beneficial for refineries to utilize existing equipment for transformation, and at least two medium-pressure hydrogenation reactors are required, each with one bed.

Claims

1. A method for producing industrial white oil by hydrogenation, comprising the following contents: (1) Paraffin-based, intermediate-based or cycloalkyl straight-run diesel is mixed with hydrogen and then enters the hydrotreating reactor, where it contacts the hydrotreating catalyst for a hydrotreating reaction; (2) The hydrotreating effluent obtained in step (1) enters a medium-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 gas containing H2S, NH3, hydrogen and part of light oil is obtained at the top of the tower; (3) The effluent gas obtained in 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 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 enters a supplementary refining reactor after pressure increase and heat exchange; (4) The tower bottom oil obtained in step (2) from which H2S and NH3 have been removed is mixed with hydrogen and then enters the supplementary refining reactor together with the cold low fraction oil from which H2S and NH3 have been removed, firstly contacts and reacts with the supplementary refining catalyst at the top, then contacts and reacts with the isomerization dewaxing catalyst at the bottom to undergo a pour point reduction reaction, and then contacts and reacts with the supplementary refining catalyst at the bottom to undergo a post-refining reaction; (5) The reaction effluent obtained in step (4) enters a fractionation system to obtain Class II industrial white oil that meets the NB / SH / T 0006-2017 standard and a small amount of a series of light white oils that meet the technical index requirements of light white oil (II) in NB / SH / T 0913-2015; in, The stripping pressure of the medium-pressure stripping tower is the same as the pressure of the hydrotreating reactor, and the pressure of the hydrotreating 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 cold high-fraction gas phase obtained in step (3) is selectively desulfurized and then re-enters the hydrotreating reactor as circulating hydrogen.

3. The method according to claim 1, characterized in that: The straight-run diesel is mixed with a catalytic cracking diesel fraction.

4. The method according to claim 1, characterized in that The straight-run diesel has an aromatic content of 1 to 99 wt % and a density of 0.825 to 0.915 g / cm at 20°C. 3 .

5. The method according to claim 1, characterized in that: The initial distillation point of the straight-run diesel is 160-200°C, and the final distillation point is 300-400°C.

6. The method according to claim 1, characterized in that The hydrotreating reactor also includes a hydrotreating protection catalyst, which is arranged upstream of the hydrotreating catalyst.

7. The method according to claim 1, characterized in that The medium-pressure stripping tower is filled with packing, and a ZnO bed layer is arranged at the lower part of the packing area.

8. The method according to claim 1 or 7, characterized in that: The stripping temperature of the medium-pressure stripping tower is 150-400° C., and the amount of hydrogen used is higher than the chemical hydrogen consumption of hydrofining.

9. The method according to claim 8, characterized in that The stripping temperature of the medium-pressure stripping tower is the outlet temperature of the hydrofining reactor, and the amount of hydrogen used is equivalent to the sum of the chemical hydrogen consumption of the hydrofining and the dissolved hydrogen of the hydrofining system.

10. The method according to claim 1, characterized in that The supplementary refining catalysts at the upper and lower parts of the supplementary refining reactor are reduced hydrorefining catalysts, and the active metals thereof are one or both of Pt and Pd, or reduced nickel.

11. The method according to claim 1, characterized in that: The isomerization dewaxing catalyst includes a carrier and an active metal, the carrier includes alumina and a 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; based on the weight of the catalyst, the content of the TON structure molecular sieve is 30wt%~80wt%, the content of the active metal component is 0.1wt%~5.0wt%, and the content of the auxiliary is 0.1wt%~5.0wt%.

12. The method according to claim 11, characterized in that The TON structure molecular sieve is selected from at least one of NU-10 molecular sieve, ZSM-22, ZSM-23 molecular sieve and ZSM-48 molecular sieve.

13. The method according to claim 12, characterized in that The TON structure molecular sieve is ZSM-22 molecular sieve.

14. The method according to claim 1, characterized in that The reaction conditions of the hydrogenation refining section in the refining reactor are: reaction temperature of 100℃~300℃, hydrogen partial pressure of 3.0MPa~18.0MPa, volume space velocity of 0.2h -1 ~3.0h -1 , hydrogen-oil volume ratio 400:1~1500:

1.

15. The method according to claim 1, characterized in that The reaction conditions of the isomerization dewaxing section in the supplementary refining reactor are: reaction temperature of 100℃~350℃, hydrogen partial pressure of 3.0MPa~18.0MPa, volume space velocity of 0.2h -1 ~5.0h -1 , hydrogen-oil volume ratio 400:1~1500:1.

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