A combined production process of processing coking liquefied gas and coking diesel

Abstract

The application discloses a combined production process of coking liquefied gas and coking diesel oil, which comprises the following contents: (1) coking diesel oil taken out from a side line of a coking device fractionating tower is cooled by heat exchange and then is introduced from the top of a first fixed bed reactor, and coking liquefied gas after removal of hydrogen sulfide is introduced from the middle lower part of the first fixed bed reactor to carry out sulfide reaction; (2) liquid phase materials flowing out from the bottom of the first fixed bed reactor are introduced into a second fixed bed reactor after hydrogen mixing to carry out gas phase hydrogenation reaction; (3) the effluent of the second fixed bed reactor is subjected to gas-liquid separation, and the separated liquid phase components are introduced into a third fixed bed reactor after pressurization to carry out liquid phase hydrogenation reaction, and the liquid phase hydrogenation reaction products are subjected to stripping and fractionation to obtain refined products. The process deeply couples the coking liquefied gas and the coking diesel oil, removes sulfides and olefins, avoids the problem that the storage of the coking diesel oil leads to the polymerization of olefins, and increases the processing difficulty, so that the processing difficulty, hydrogen consumption and energy consumption of the device are reduced.

Description

Technical Field

[0001] This invention belongs to the field of clean oil refining, and specifically relates to a combined production process for processing coking liquefied gas and coking diesel. Background Technology

[0002] Delayed coking processes utilize heavy oil products, such as vacuum residue and cracked residue, as feedstock. Under high-temperature conditions, thermal cracking and condensation reactions occur to produce components such as gas, gasoline, diesel, wax oil, and petroleum coke. However, due to the poor properties and high density of the feedstock in coking units, which contains significant amounts of sulfides, nitrides, and other impurities, coupled with the inherent characteristics of coking—primarily thermal cracking—the products, such as coking liquefied gas and coking diesel, contain substantial amounts of olefins. Olefins, especially dienes, are highly susceptible to oxidation, leading to the formation of gums or coking precursors. This results in unstable oil properties and poor stability. Generally, hydrorefining is required to remove nitrogen-, sulfur-, and olefin-containing components before the oil can be used as a blending component for automotive diesel or as a chemical feedstock. Coking diesel contains a large amount of olefins and dienes. If stored for too long, the olefins, especially dienes, in the oil are prone to condensation reactions, which leads to the formation of more coking precursors such as gums in the oil, significantly affecting the processing difficulty and long-term operation of the oil.

[0003] Coking liquefied petroleum gas (LPG) contains not only a significant amount of acidic components such as H2S, COS, and CS2, but also a large quantity of sulfides, such as methanethiol (CH3SH) and ethanethiol (C2H5SH). During combustion, these sulfides produce harmful SOx emissions. Furthermore, sulfides like mercaptans can easily poison downstream catalysts and promote the oxidation of reactive hydrocarbons in the LPG, forming gum-like substances. Currently, coking liquefied petroleum gas and coking diesel are generally processed separately, increasing refinery equipment investment and floor space requirements.

[0004] CN106190234A discloses a hydrodesulfurization and refining process for coking diesel. The process employs a fixed-bed reactor loaded with a catalyst active for hydrodesulfurization and denitrification. The active component of this type of catalyst is a nitride or carbide metal, wherein the metal component is molybdenum, tungsten, etc. It can operate at 300–450°C, a hydrogen partial pressure of 6.0–9.0 MPa, and a volume hourly space velocity of 1.5–3 h⁻¹. -1 Under a hydrogen-to-oil volume ratio of 500, the sulfide content in the feedstock is controlled to below 5 ppm. However, this process is complex, requiring the feedstock to pass through a tank area and undergo pre-fractionation before entering the feedstock buffer tank of the reaction unit for further reaction. The equipment is extensive and requires significant investment.

[0005] CN101003745A discloses a method for producing high-quality diesel from coking full-fraction oil. The method involves separating the coking full-fraction oil into light and heavy fractions, which are then subjected to hydrorefining and hydromodification, respectively. The reaction products from both processes are then combined in a fractionation system to separate naphtha and diesel. This method has a relatively complex reaction process. When the heavy fraction enters the hydromodification reaction zone, the nitrogen compounds within it can severely affect the activity of the modifier.

[0006] US2882224 discloses a method for removing mercaptan from liquefied petroleum gas (LPG). The process involves contacting LPG with a sodium hydroxide alkaline solution containing a sulfonated cobalt phthalocyanine or polycobalt phthalocyanine catalyst in an extraction tower to remove mercaptan. The sodium thiolate is then transferred to the alkaline solution, which enters a regeneration tower. In the tower, the sodium thiolate is converted to disulfides by the sulfonated cobalt phthalocyanine catalyst and air. The disulfides are then separated from the alkaline solution, which is then recycled. However, the sulfonated cobalt phthalocyanine catalyst used in this process is prone to aggregation and deactivation in the alkaline environment, resulting in a short alkaline solution circulation cycle and a large amount of waste alkaline emissions. Furthermore, during mercaptan extraction, dissolved oxygen in the alkaline solution, under the influence of the catalyst, can cause some sodium thiolate to be converted to disulfides and introduced into the refined LPG, leading to excessive total sulfur content in the LPG.

[0007] CN104194833A discloses a deep desulfurization process for liquefied petroleum gas (LPG). The process involves passing LPG through an amine extraction tower to remove hydrogen sulfide, followed by a sulfur transfer reactor to remove sulfides. The LPG then enters a stabilization unit where it is separated from high-boiling-point sulfides by distillation to obtain ultra-low sulfur LPG. This desulfurization process is relatively complex, with the reaction and separation occurring in two separate units. Summary of the Invention

[0008] To address the shortcomings of existing technologies, this invention provides a combined production process for processing coking liquefied petroleum gas (LPG) and coking diesel. This process deeply couples the two by controlling reaction conditions, removing sulfides and olefins. This process can also directly process side-stream feedstocks from coking units, avoiding the problem of increased gum content and production difficulty caused by the polymerization of olefins, especially dienes, during the storage of coking diesel. This reduces the processing difficulty, hydrogen consumption, and energy consumption of the unit.

[0009] The combined production process of coking liquefied petroleum gas and coking diesel of the present invention includes the following:

[0010] (1) The coking diesel fuel extracted from the side stream of the fractionation tower of the coking unit is cooled by heat exchange and then enters from the top of the first fixed bed reactor. The coking liquefied gas after removing hydrogen sulfide enters from the lower part of the first fixed bed reactor. The coking diesel fuel and coking liquefied gas undergo a sulfurization reaction in the catalyst bed set in the upper part of the reactor. The gaseous material after the reaction flows out from the top side stream of the reactor, and the liquid material flows out from the bottom of the reactor.

[0011] (2) The liquid material flowing out of the bottom of the first fixed-bed reactor enters the second fixed-bed reactor after being mixed with hydrogen, and carries out gas-phase hydrogenation desulfurization and deolefination reaction.

[0012] (3) The effluent from the second fixed-bed reactor is subjected to gas-liquid separation. The separated gas phase component is condensed to obtain light naphtha component. The separated liquid phase component is pressurized and enters the third fixed-bed reactor for liquid-phase hydrogenation reaction. The liquid-phase hydrogenation reaction product is stripped and fractionated to obtain refined product.

[0013] In the process of this invention, the liquefied petroleum gas (LPG) after hydrogen sulfide removal has an olefin content of 5%–60 vol%, preferably ≤50 vol%, and a mercaptan sulfur content of 20–500 μg / g, preferably 50–200 μg / g. Generally, the main components of coking LPG from coking plants are C2, C3, and C4 components, along with a certain amount of olefins, hydrogen sulfide, mercaptan sulfur, etc. The amount of mercaptan sulfur and olefins in the LPG is related to the type of delayed coking unit, raw materials, and reaction conditions. The main sulfide in the LPG is hydrogen sulfide, which accounts for more than 90% of the total sulfur in the LPG. Organic sulfur mainly consists of methanethiol, ethanethiol, and small amounts of carbonyl sulfides.

[0014] In the process of this invention, the properties of the coking diesel oil are as follows: distillation range of 160–360℃, density ≤ 0.88 g / m³. 3 , S≯15000μg / g, N≯1000μg / g.

[0015] In the process of this invention, the temperature of the coking diesel oil extracted from the side stream of the fractionation tower of the coking unit is generally 250-350°C. It needs to be heat exchanged before entering the first fixed bed reactor. It can generally exchange heat with the effluent of the first reactor. After the heat exchange, corresponding cooling equipment is set up to control the inlet temperature of the first reactor at 60-200°C.

[0016] In this invention, the upper part of the first fixed-bed reactor is filled with a sulfideation catalyst, such as a MoCo-type catalyst or a Ni-based catalyst. The primary reaction in the first fixed-bed reactor involves the sulfideation of thiols from coking liquefied petroleum gas (LPG) with olefins from LPG and coking diesel, producing heavy sulfides and thiols. These heavy sulfides and thiols, along with the coking diesel, flow out as liquid phase material from the bottom of the first fixed-bed reactor. The gaseous material after the reaction is separated by a separator to obtain dry gas and desulfurized LPG. The operating conditions of the first fixed-bed reactor are: reaction pressure 0.1–3.0 MPa, preferably 0.5–2.0 MPa; reaction temperature 50–300°C, preferably 80–180°C; and liquid hourly space velocity (LHSV) 0.1–20 h⁻¹. -1 Preferably 1-10h -1 .

[0017] In the process of this invention, the amount of hydrogen in the hydrogen mixing process described in step (2) depends on the amount of sulfur, nitrogen, and olefins in the extracted oil, as well as the reaction conditions. The liquid phase material before hydrogen mixing is generally heated to meet the reaction temperature of the second fixed-bed reactor through heat exchange or heating.

[0018] In the process of this invention, the gas-phase hydrogenation reaction mainly involves the liquid-phase material from the first fixed-bed reactor undergoing a hydrogenation refining reaction in the gas phase within the second fixed-bed reactor. By loading a catalyst with strong desulfurization performance into the second fixed-bed reactor, the remaining olefins and most of the sulfides in the coking diesel can be removed under high temperature, low pressure, and high space velocity conditions. The second fixed-bed reactor has a higher reaction temperature and a lower reaction pressure, which is conducive to the desulfurization reaction. The higher reaction temperature is not conducive to the saturation reaction of aromatics, and it can target the removal of different components in the reactants.

[0019] The catalyst packed in the second fixed-bed reactor includes a protective agent, a silicon scavenger, and a hydrorefining catalyst. The protective agent, silicon scavenger, and hydrorefining catalyst comprise a support and a hydrorefining active metal; wherein the support is an inorganic refractory oxide selected from one or more of alumina, amorphous aluminum silicate, silica, or titanium dioxide; the hydrorefining active metal comprises Group VIB and / or Group VIII metal components, wherein Group VIB is selected from tungsten and / or molybdenum, and Group VIII is selected from nickel and / or cobalt. The protective agent is, for example, the FBN Bird's Nest series protective agent developed by Sinopec (Dalian) Petrochemical Research Institute Co., Ltd. (FRIPP); the silicon scavenger is, for example, the FHRS series silicon scavenger developed by FRIPP; and the hydrorefining catalyst is, for example, the FHUDS series catalyst developed by FRIPP.

[0020] In the process of this invention, the gas-phase hydrogenation reaction conditions in the second fixed-bed reactor are as follows: pressure 0.1–3.0 MPa, preferably 0.5–2.0 MPa; hydrogen-to-oil volume ratio 20–600, preferably 50–200; temperature 250–400°C, preferably 300–380°C; volume hourly space velocity (VHSV) 0.1–6.0 h⁻¹. -1 Preferably 1.0 to 4.0 h -1 .

[0021] In this invention, the gaseous components separated by the separator are condensed to obtain light naphtha components along with hydrogen, ammonia, and hydrogen sulfide. The light naphtha components can be used as feedstock for ethylene cracking. After separation, the hydrogen sulfide and other substances generated after the second reactor's reaction to remove sulfides are separated at the top of the separator, ensuring that the material entering the third reactor is free of hydrogen sulfide. This avoids the problem of hydrogen sulfide in the material competing for adsorption with subsequent deep hydrotreating and deaeration reactions, which would increase the difficulty of the reaction. Simultaneously, due to the characteristics of hydraulic decoking in coking units, the coking oil is prone to contain a certain amount of water. A separator is installed between the second and third fixed-bed reactors to remove water from the reaction system, ensuring that the material entering the third fixed-bed reactor is free of moisture. This helps maintain the stability of catalyst strength and activity, and is beneficial for the reaction to proceed smoothly.

[0022] In this invention, the separated liquid phase components are pressurized by a compressor, mixed with hydrogen, and then enter the liquid phase reactor. The compressor can be a conventional commercial compressor, such as a reciprocating or centrifugal compressor. The hydrogen mixed in the hydrogen mixer is mainly to meet the reaction conditions of the third fixed-bed reactor, providing the required hydrogen for the liquid phase hydrogenation reaction to carry out reactions such as hydrodearomatization and hydrodenitrification. The hydrogen mixer can be a common commercial baffle-type hydrogen mixer or other types of hydrogen mixers.

[0023] In the process of this invention, the liquid-phase reaction conditions in the third fixed-bed reactor are as follows: pressure 0.5–12.0 MPa, preferably 4–10.0 MPa; temperature 200–400 °C, preferably 300–380 °C; volume hourly space velocity (VHSV) 0.1–4.0 h⁻¹. -1 Preferably 0.5–2.0 h -1 .

[0024] In the process of this invention, the third fixed-bed reactor is filled with a hydrorefining catalyst, which includes a support and a hydrorefining active metal component. The support is an inorganic refractory oxide, selected from one or more of alumina, amorphous aluminum silicate, silica, or titanium dioxide. The hydrorefining active metal includes a Group VIII metal component, with Group VIII metals selected from platinum and / or palladium. The hydrorefining catalyst is, for example, the FHUDS series catalysts developed by FRIPP. This process removes nitrides and aromatics from the material.

[0025] Furthermore, the liquid-phase hydrogenation reaction product is mixed with light naphtha components and then enters the stripping and fractionation system for stripping and fractionation.

[0026] Compared with the prior art, the method of the present invention has the following advantages:

[0027] (1) By controlling the reaction conditions of the first fixed-bed reactor, the present invention reacts the small molecule sulfides in coking liquefied gas with the olefins in liquefied gas and / or coking diesel, and transforms the reaction products into high-boiling-point, large-molecule thiols, thioethers and other substances. At the same time, most of the olefin components in the liquefied gas are retained from saturation reaction, and the high-boiling-point thiols and thioethers generated are directly carried out of the reaction system by coking diesel for subsequent gas-phase hydrogenation reactor to carry out hydrodesulfurization and olefin saturation reaction.

[0028] (2) This invention can also utilize the relatively reactive diene components in coking diesel to carry out a thioetherification reaction with thiols in liquefied petroleum gas. This not only removes dienes from the oil but also pre-treats the coking diesel, avoiding the problem of dienes coking in heat exchangers, furnaces, reactors, etc., when coking diesel is processed alone. In addition, the introduction of coking diesel can provide more olefins for the thioetherification reaction, which is beneficial for the removal of thiols and sulfur substances in the reactor. At the same time, the introduced coking diesel can also act as a stabilizer in this system, carrying away the large molecular sulfides generated after the thioetherification reaction, which is conducive to the forward progress of the thioetherification reaction.

[0029] (3) This invention can be directly coupled with the fractionation tower of the coking unit. By making reasonable use of the heat exchange of the diesel fraction on the side line of the fractionation tower of the coking unit, the directly extracted coking diesel fraction can be put into the first fixed bed reactor. The coking diesel feedstock can enter the hydrogenation reactor directly without passing through the tank area or buffer tank. This can greatly reduce the amount of diene and olefin gum formed in the feedstock due to the long storage time of coking diesel, thereby delaying the increase of bed pressure drop and the amount of carbon deposits in the catalyst in the unit, and thus extending the operating cycle of the unit.

[0030] (4) By installing a separator between the second and third fixed-bed reactors, this invention can remove hydrogen sulfide generated in the second fixed-bed reactor, ensuring that the material entering the third fixed-bed reactor is free of hydrogen sulfide. This avoids hydrogen sulfide competing with aromatics and nitrogen compounds for adsorption on the catalyst in the third fixed-bed reactor, which is beneficial for the hydrodearomatization and denitrification reactions. Simultaneously, the separator can also separate water present in the oil and water that may be generated during the reaction, ensuring that the water entering the third reactor does not affect anhydrous substances. This prevents the catalyst strength from being damaged in the presence of high water content, affecting the catalyst's activity and stability. Furthermore, it reduces the partial pressure of water in the gas phase, which is beneficial for the hydrotreating reaction.

[0031] (5) This invention organically combines gas-phase hydrogenation and liquid-phase hydrogenation reactors, setting different reaction conditions based on the advantages of different reactors, allowing the desulfurization reaction and the denitrification and dearomatization reactions to occur separately, thus reducing the hydrogen consumption of the unit. The second fixed-bed reactor is a gas-phase hydrogenation reactor, which, by loading a catalyst with high desulfurization activity, can remove most of the sulfides in the feedstock under high temperature, low pressure, and a large hydrogen-to-oil ratio, and saturate the olefins. At the same time, due to the higher temperature and lower pressure, the denitrification and dearomatization reactions are suppressed, which can also avoid the problem of mutual inhibition between desulfurization and denitrification reactions, thereby improving the reaction efficiency. The third fixed-bed reactor is a liquid-phase hydrogenation reaction, which is conducive to the denitrification and dearomatization reactions under higher pressure conditions. At the same time, hydrogen is added at the reactor inlet and in the middle of the bed to ensure the hydrogen required for the hydrogenation reaction of the unit. Separators are set in the two reactors to remove small molecules and hydrogen sulfide from the material, which is conducive to the unit maintaining the liquid phase. At the same time, the main components entering the third fixed-bed reactor are large molecular polycyclic aromatic hydrocarbons and large molecular sulfides, thereby improving the conversion rate of aromatic hydrogenation reaction. Attached Figure Description

[0032] Figure 1 This is a schematic diagram of a combined production process for processing coking liquefied gas and coking diesel according to the present invention.

[0033] Among them, 1-desulfurized coking liquefied gas, 2-coking diesel, 3-hydrogen, 4-first fixed bed reactor, 5-liquid separator, 6-dry gas, 7-liquefied gas, 8-heat exchanger, 9-hydrogen mixer, 10-second fixed bed reactor, 11-separator, 12-gas-liquid separator, 13-high-grade gas, 14-light naphtha component, 15-booster pump, 16-hydrogen mixer, 17-third fixed bed reactor, 18-diesel component. Detailed Implementation

[0034] The present invention will be further described below with reference to embodiments, but it should be understood that the scope of protection of the present invention is not limited to the embodiments. In the present invention, unless otherwise expressly stated, percentages and contents are all expressed by mass.

[0035] The following is combined with Figure 1 The process flow of the present invention will be described in detail.

[0036] After desulfurization, the coking liquefied gas component 1 is mixed with hydrogen 3 and fed from the bottom of the first fixed-bed reactor 4. The coking diesel 2 is fed from the top of the first fixed-bed reactor 4. A sulfur etherification reaction occurs in the first fixed-bed reactor 4. The gaseous products are extracted from the upper side stream of the first fixed-bed reactor 4, pass through the separator 5, and are separated into dry gas 6 and liquefied gas 7. The liquid component flows out from the bottom of the first fixed-bed reactor 4, passes through the heat exchanger 8, and enters the hydrogen mixer 9 to mix with hydrogen 3 before entering the second fixed-bed reactor 10 for a hydrogenation reaction. The reactor material enters the high-efficiency separator 11, and enters the separator tank 12 to separate high-efficiency gas 13 and light naphtha 14. The liquid component separated from the high-efficiency separator is pressurized by the booster pump 15 and enters the hydrogen mixer 16 to mix with hydrogen 3 before entering the third fixed-bed reactor 17 for a deep hydrorefining reaction. The diesel component 18 after the reaction enters the subsequent gas stripping fractionation equipment to obtain diesel and naphtha components.

[0037] Examples 1-3

[0038] The first fixed-bed reactor is loaded with 50 mL of sulfideation catalyst C, with alumina-silica support. The active components are Mo and Co, with the support content being 79%, MoO3 content being 16%, CoO content being 3%, and a small amount of additives. The second fixed-bed reactor is loaded with 50 mL of Mo-Co type catalyst A, with alumina support. The active metal components are Mo and Co, with the support content being 74%, Mo (based on oxide content) being 23%, Co (based on oxide content) being 5%, and the remainder being support. The third fixed-bed reactor is loaded with 50 mL of Mo-Ni type catalyst B, with Mo and Ni active metal components, and a small amount of additives. The active metal content is 24% (based on oxide content), Ni (based on oxide content) being 5%, and a small amount of additives, with the remainder being support. In the first reactor, catalyst C is in the upper part. Liquefied gas and hydrogen are fed from the lower part of the reactor, and coking diesel is fed from the upper part of the first reactor. Liquefied gas product is extracted from the upper side stream of the first reactor and the product is separated by a separator. A separator and a corresponding separator are set between the second and third fixed-bed reactors. A stripping fractionation device is set after the third reactor.

[0039] The catalyst properties are shown in Table 1, the properties of the coking diesel feedstock are shown in Table 2, the desulfurized coking liquefied gas is shown in Table 3, and the reaction process conditions and evaluation results are shown in Table 4.

[0040] Comparative Example 1

[0041] A single conventional fixed-bed hydrogenation reactor was used. The reactor was loaded with 100 mL of Mo-Co type catalyst A, and a separator and stripping process were installed downstream of the reactor. Hydrogen was passed through in a single pass. The diesel feedstock was the same as in Examples 1-3, and the reaction conditions and results are shown in Table 3.

[0042] Comparative Example 2

[0043] A conventional fixed-bed hydrogenation reactor was used, and the reaction was carried out under normal conditions. The reactor was loaded sequentially from top to bottom with 60 mL of Mo-Ni type catalyst B and 40 mL of Mo-Co type catalyst A (same as in Examples 1-3). A separator and stripping process were installed downstream of the reactor as usual. Hydrogen was passed through in a single pass. The diesel feedstock was the same as in Examples 1-3. The reaction process conditions and results are shown in Table 3.

[0044] Comparative Example 3

[0045] A diesel liquid-phase hydrogenation process was adopted, and a hydrogenation reactor was set up. The reactor was loaded sequentially from top to bottom with 60 mL of Mo-Ni type catalyst B and 40 mL of Mo-Co type catalyst A (same as in Examples 1-3). The high-precision, low-precision, and stripping processes were set up normally after the reactor. The diesel feedstock was the same as in Examples 1-3, and the reaction process conditions and results are shown in Table 3.

[0046] Comparative Example 4

[0047] A conventional fixed-bed reactor is used, with 50 mL of sulfide etherification catalyst C inside. After desulfurization, the coking liquefied gas and hydrogen enter the reactor from the top.

[0048] Table 1. Physicochemical properties of catalysts

[0049] Catalyst number Thioetherifying agent C A B Active metals Mo-Ni Mo-Co Mo-Ni <![CDATA[MoO3,wt%]]> 16 23 24 CoO, wt% 3 5 - NiO, wt% - - 5 shape cylindrical bar Clover Clover Diameter, mm 1.1～1.3 1.1～1.3 1.1～1.3 <![CDATA[Specific surface area, m 2 ·g -1 > 268 215 202 <![CDATA[Pore volume, mL·g -1 > - 0.35 0.34

[0050] Table 2 Properties of Coking Diesel Feedstock

[0051] Oil properties Analyze data <![CDATA[Density (20 °C), kg·m -3 > 856.8 Distillation range, ℃ 165～365 Sulfur content, μg / g 11640 Total aromatics, wt% 27.8 Polycyclic aromatic hydrocarbons (wt%) 12.8

[0052] Table 3 Properties of Liquefied Petroleum Gas Feedstock

[0053] Properties of liquefied gas Analyze data Olefin content, v% 24.2 Total S, μg / g 168 Thiol sulfur 84

[0054] Table 3 Hydrogenation process conditions and results

[0055]

[0056]

Claims

1. A combined production process for processing coking liquefied gas and coking diesel, comprising the following: (1) Coking diesel extracted from the side stream of the fractionation tower of the coking unit is cooled by heat exchange and enters from the top of the first fixed bed reactor; coking liquefied gas after removing hydrogen sulfide enters from the lower part of the first fixed bed reactor; coking diesel and coking liquefied gas undergo a sulfurization reaction in the catalyst bed set in the upper part of the reactor; the gaseous material after the reaction flows out from the top side stream of the reactor, and the liquid material flows out from the bottom of the reactor; (2) The liquid material flowing out from the bottom of the first fixed bed reactor is subjected to... After passing through hydrogen, it enters the second fixed-bed reactor for gas-phase hydrogenation desulfurization and deolefination reaction; (3) the effluent from the second fixed-bed reactor is subjected to gas-liquid separation, the separated gas phase component is condensed to obtain light naphtha component, the separated liquid phase component is pressurized and enters the third fixed-bed reactor for liquid-phase hydrogenation reaction, the liquid-phase hydrogenation reaction product is stripped and fractionated to obtain refined product; the operating conditions of the first fixed-bed reactor are: reaction pressure 0.1~3.0MPa, reaction temperature 50~300℃, liquid hourly space velocity 0.1~20h -1 The gas-phase hydrogenation reaction conditions in the second fixed-bed reactor are as follows: pressure 0.1~3.0 MPa, hydrogen-to-oil volume ratio 20~600, temperature 250~400℃, and volume hourly space velocity 0.1~6.0 h⁻¹. -1 The liquid-phase reaction conditions in the third fixed-bed reactor are as follows: pressure 4~12.0 MPa, temperature 200~400℃, and volume hourly space velocity 0.1~4.0 h⁻¹. -1 .

2. The process according to claim 1, characterized in that: The liquefied coke gas after hydrogen sulfide removal has an olefin content of 5v%~60v% and a mercaptan sulfur content of 20~500μg / g.

3. The process according to claim 1, characterized in that: The properties of the coking diesel oil are as follows: distillation range of 160~360℃, density ≤0.88g / m³. 3 , S≯15000μg / g, N≯1000μg / g.

4. The process according to claim 1, characterized in that: The temperature of the coking diesel oil drawn from the side stream of the fractionation tower of the coking unit is 250~350℃. After heat exchange, it enters the first fixed bed reactor, and the inlet temperature of the first fixed bed reactor is controlled at 60~200℃.

5. The process according to claim 1, characterized in that: The upper part of the first fixed-bed reactor is filled with a sulfideation catalyst; the operating conditions of the first fixed-bed reactor are: reaction pressure of 0.5~2.0 MPa, reaction temperature of 80~180℃, and liquid hourly space velocity of 1~10 h⁻¹. -1 .

6. The process according to claim 1, characterized in that: The catalyst packed in the second fixed-bed reactor includes a hydrogenation protectant, a silicon scavenger, and a hydrogenation refining catalyst. The gas-phase hydrogenation reaction conditions within the second fixed-bed reactor are as follows: pressure 0.5–2.0 MPa, hydrogen-to-oil volume ratio 50–200, temperature 300–380 °C, and volume hourly space velocity (VHSV) 1.0–4.0 h⁻¹. -1 .

7. The process according to claim 1, characterized in that: The gas phase components separated in step (3) are condensed to obtain light naphtha components, hydrogen, ammonia and hydrogen sulfide. The light naphtha components are used as ethylene cracking feedstock. After separation, the hydrogen sulfide generated after the second fixed bed reactor reacts to remove sulfides is separated from the reaction system at the top of the separator.

8. The process according to claim 1, characterized in that: The third fixed-bed reactor is filled with a hydrorefining catalyst, which includes a support and a hydrorefining active metal component. The carrier is an inorganic refractory oxide selected from one or more of alumina, amorphous aluminum silica, silicon dioxide, or titanium dioxide; the hydrogenated active metal includes a Group VIII metal component, with Group VIII selected from platinum and / or palladium.

9. The process according to claim 1, characterized in that: The liquid-phase hydrogenation reaction product is mixed with light naphtha components and then enters the stripping and fractionation system for stripping and fractionation.

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