Joint production process for processing coked liquefied gas and coker diesel oil

Abstract

The invention discloses a joint production process for processing coked liquefied gas and coker diesel oil, which comprises the following steps: (1) coker diesel oil extracted from a side line of a fractionating tower of a coking device enters from the top of a first fixed bed reactor after being subjected to heat exchange and cooling; the coking liquefied gas with hydrogen sulfide removed enters the first fixed bed reactor from the middle lower part of the first fixed bed reactor for a thioetherification reaction; (2) a liquid-phase material flowing out of the bottom of the first fixed bed reactor is mixed with hydrogen and then enters a second fixed bed reactor for a gas-phase hydrogenation reaction; and (3) carrying out gas-liquid separation on the effluent of the second fixed bed reactor, pressurizing the separated liquid-phase component, feeding the pressurized liquid-phase component into a third fixed bed reactor, carrying out a liquid-phase hydrogenation reaction, and carrying out steam stripping and fractionation on the liquid-phase hydrogenation reaction product to obtain a refined product. According to the process, the two are deeply coupled to remove sulfides and olefins, so that the problems of olefin polymerization and processing difficulty increase caused by coking diesel storage are avoided, and the processing difficulty, hydrogen consumption and energy consumption of the device are reduced.

Description

Technical Field

[0001] The invention belongs to the field of clean oil refining, and in particular relates to a combined production process for processing coking liquefied gas and coking diesel. Background Art

[0002] The delayed coking process can use heavy oil products, such as vacuum residue and cracked residue, as raw materials, and can carry out thermal cracking and condensation reactions under high temperature conditions to produce components such as gas, gasoline, diesel, wax oil and petroleum coke. Due to the poor properties and high density of the raw materials of the coking unit, and the presence of more impurities such as sulfides and nitrides, coupled with the characteristics of the coking reaction itself, thermal cracking reactions mainly occur, resulting in more olefins in the coking reaction products such as coking liquefied gas and coking diesel. Olefin substances, especially dienes, are very easy to oxidize, and then generate colloids or coking precursors, resulting in unstable oil properties and poor stability. Generally, they need to be hydrotreated to remove nitrogen, sulfur and olefin components before they can be used as automotive diesel blending components or chemical raw materials. Coking diesel contains a large amount of olefins and diolefins. If it is stored for too long, the olefins, especially diolefins, in the oil are prone to condensation reactions, which will lead to the formation of more gums and other coking precursors in the oil, which will greatly affect the processing difficulty and long-term operation of the oil.

[0003] In addition to containing more acidic components such as H2S, COS, CS2, coking liquefied gas also contains a large amount of sulfides, such as methyl mercaptan (CH3SH) and ethyl mercaptan (C2H5SH), which will produce harmful SOx emissions during the combustion process. Sulfides such as mercaptans can also easily poison downstream catalysts and promote the oxidation reaction of active hydrocarbons in liquefied gas to form colloids. At present, coking liquefied gas and coking diesel are generally processed separately, which increases the equipment investment and floor space of refineries.

[0004] CN106190234A discloses a coking diesel oil hydrodesulfurization refining process, which uses a fixed bed reactor, in which a catalyst with hydrodesulfurization and denitrification activity is loaded, and the active components of this type of catalyst are nitrided and / or carburized metals, wherein the metal components are molybdenum, tungsten, etc., and can be heated to 300-450°C, with a hydrogen partial pressure of 6.0-9.0 MPa and a volume space velocity of 1.5-3h -1 , under the condition of hydrogen-to-oil volume ratio of 500, the sulfide in the raw material is controlled below 5ppm. However, the process is complicated, the raw material needs to pass through the tank area and then undergo pre-fractionation before entering the raw material buffer tank of the reaction unit for reaction, and the equipment is large and the investment is large.

[0005] CN101003745A discloses a method for producing high-quality diesel from coking whole fraction oil, wherein the coking whole fraction oil is separated into light fraction and heavy fraction, and then subjected to hydrofining and hydromodification, respectively, and the reaction products of the two are then separated into naphtha and diesel by entering a fractionation system together. The reaction process of this method is relatively complicated, and the heavy fraction oil enters the hydromodification reaction zone, where the nitrides will have a serious impact on the activity of the modifier.

[0006] US2882224 discloses a method for removing mercaptans from liquefied petroleum gas. The technical process is as follows: sodium hydroxide lye containing sulfonated cobalt phthalocyanine or polyphthalocyanine cobalt catalyst is fully contacted with liquefied petroleum gas in an extraction tower to remove mercaptans from the liquefied gas, and sodium mercaptide is transferred to the lye. The lye containing sodium mercaptide enters a regeneration tower, and sodium mercaptide is converted into disulfide under the action of the sulfonated cobalt phthalocyanine catalyst and air in the tower, and the disulfide and lye are separated, and the lye is recycled after regeneration. However, the sulfonated cobalt phthalocyanine catalyst used in this process is easily aggregated and deactivated in alkali, resulting in a short lye cycle and a large amount of waste alkali discharge; when the lye containing the catalyst is used for mercaptan extraction, part of the sodium mercaptide is easily converted into disulfide under the action of the dissolved oxygen in the lye and the catalyst and brought into the refined liquefied petroleum gas, causing the total sulfur content of the petroleum gas to exceed the standard.

[0007] CN104194833A discloses a method for deep desulfurization of liquefied gas. The process flow is that the liquefied gas passes through an amine liquid extraction tower to remove hydrogen sulfide, then enters a sulfur transfer reactor to remove sulfides in the liquefied gas, and then enters a stabilization unit to separate the liquefied gas from high-boiling-point sulfides by distillation to obtain ultra-low sulfur liquefied gas. The liquefied gas desulfurization process is relatively complex, and the reaction separation is carried out in two units separately. Summary of the invention

[0008] In view of the shortcomings of the prior art, the present invention provides a joint production process for processing coking liquefied gas and coking diesel. The process deeply couples the two by controlling the reaction conditions to remove sulfides and olefins therein. The process can also directly process side-line raw materials from the coking unit, thereby avoiding the problem of increased colloids due to the polymerization of olefins, especially diolefins, caused by the storage of coking diesel, which increases the difficulty of production and processing, and reduces the processing difficulty, hydrogen consumption and energy consumption of the device.

[0009] The combined production process of processed coking liquefied gas and coking diesel oil of the present invention comprises the following contents:

[0010] (1) After heat exchange and cooling, the coking diesel extracted from the side line of the fractionation tower of the coking unit enters from the top of the first fixed bed reactor, and the coking liquefied gas after the removal of hydrogen sulfide enters from the middle and lower part of the first fixed bed reactor. The coking diesel and the coking liquefied gas undergo thioetherification reaction in the catalyst bed layer arranged in the middle and upper part of the reactor. The gaseous materials after the reaction flow out from the side line at the top of the reactor, and the liquid materials flow out from the bottom of the reactor;

[0011] (2) the liquid phase material flowing out from the bottom of the first fixed bed reactor enters the second fixed bed reactor after being mixed with hydrogen to carry out gas phase hydrodesulfurization and deolefination reaction;

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

[0013] In the process of the present invention, the liquefied gas after the removal of hydrogen sulfide has an olefin content of 5% to 60v%, preferably an olefin content of ≯50v%, and a mercaptan sulfur content of 20 to 500μg / g, preferably 50 to 200μg / g. Generally, the main components of the coking liquefied gas from the coking coking unit are C2, C3, and C4 components, and there are also a certain amount of olefins and hydrogen sulfide, mercaptan sulfur and other substances, and the amount of mercaptan sulfur and olefins in the liquefied gas is related to the type of delayed coking unit, raw materials and reaction conditions. The main sulfide in the liquefied gas is hydrogen sulfide, of which hydrogen sulfide accounts for more than 90% of the total sulfur in the liquefied gas, and the organic sulfur is mainly methyl mercaptan, ethyl mercaptan and a small amount of carbonyl sulfide and other sulfides.

[0014] In the process of the present invention, the properties of the coking diesel are as follows: distillation range is 160-360°C, density is ≯0.88 g / m 3 , S≯15000μg / g, N≯1000μg / g.

[0015] In the process of the present invention, the temperature of the coking diesel extracted from the side line of the distillation tower of the coking unit is generally 250-350°C, and it needs to enter the first fixed bed reactor after heat exchange, and can generally be heat exchanged with the effluent of the first reactor. After the heat exchange, the corresponding cooling equipment is set, and the inlet temperature of the first reactor can be controlled at 60-200°C.

[0016] In the process of the present invention, the upper part of the first fixed bed reactor is filled with a thioetherification catalyst, such as a MoCo-type catalyst, a Ni-based catalyst, etc. Among them, the first fixed reactor mainly contains mercaptans in the coking liquefied gas, olefins in the liquefied gas, and olefins in the coking diesel undergo thioetherification reactions to generate heavy sulfides and mercaptans. The generated heavy sulfides and mercaptans flow out of the bottom of the first fixed bed reactor together with the coking diesel as liquid phase materials. The gas phase materials after the reaction are separated in a liquid separator to obtain dry gas and desulfurized liquefied gas. The operating conditions of the first fixed bed reactor are: reaction pressure of 0.1-3.0MPa, preferably 0.5-2.0MPa, reaction temperature of 50-300°C, preferably 80-180°C, liquid hourly volume space velocity of 0.1-20h -1 , preferably 1 to 10 hours -1 .

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

[0018] In the process of the present invention, the gas phase hydrogenation reaction is mainly a hydrogenation refining reaction of the liquid phase material of the first fixed bed reactor in the second fixed bed reactor in the form of gas phase. By loading a catalyst with strong desulfurization performance in 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, while a higher reaction temperature is not conducive to the aromatic saturation reaction, and different components in the reaction materials can be removed in a targeted manner.

[0019] The catalyst loaded 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 include a carrier and a hydrogenation active metal; wherein the carrier is an inorganic refractory oxide, selected from one or more of alumina, amorphous silica-alumina, silicon dioxide or titanium oxide; the hydrogenation active metal includes a metal component of Group VIB and / or Group VIII, wherein Group VIB is selected from tungsten and / or molybdenum, and Group VIII is selected from nickel and / or cobalt. The protective agent is such as the FBN Bird's Nest series protective agent developed by Sinopec (Dalian) Petrochemical Research Institute Co., Ltd. (FRIPP for short), the silicon scavenger is such as the FHRS series silicon scavenger developed by FRIPP, and the hydrorefining catalyst is such as the FHUDS series catalyst developed by FRIPP.

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

[0021] In the process of the present invention, the gas phase components separated by the separator are condensed to obtain light naphtha components and hydrogen, ammonia and hydrogen sulfide, etc., wherein the light naphtha components can be used as ethylene cracking materials. After separation, the second reactor can react to remove sulfides and other substances generated, and the reaction system is separated at the top of the separator to ensure that the materials entering the third reactor do not contain hydrogen sulfide substances, so as to avoid the competitive adsorption of hydrogen sulfide in the materials for subsequent deep hydrogenation, dearomatization and denitrogenation reactions, which leads to the problem of increased reaction difficulty. At the same time, due to the characteristics of hydraulic decoking of the coking unit, it is easy to cause a certain amount of water in the coking oil products. The separator is arranged between the second fixed bed reactor and the third fixed bed reactor, which can remove water from the reaction system and ensure that the materials entering the third fixed bed reactor do not contain water, which is beneficial to ensure the stability of the catalyst strength and activity and is beneficial to the reaction.

[0022] In the process of the present 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 by the hydrogen mixer is mainly used to meet the reaction conditions of the third fixed bed reactor, and to provide the required hydrogen for the liquid phase hydrogenation reaction to carry out hydrogenation dearomatization, hydrogenation denitrogenation and other reactions. The hydrogen mixer can be a common commercial baffled hydrogen mixer or other types of hydrogen mixers.

[0023] In the process of the present 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 space velocity 0.1-4.0 h -1 , preferably 0.5~2.0h -1 .

[0024] In the process of the present invention, the hydrorefining catalyst loaded in the third fixed bed reactor comprises a carrier and a hydrogenation active metal component; wherein the carrier is an inorganic refractory oxide selected from one or more of alumina, amorphous silica-alumina, silicon dioxide or titanium oxide; the hydrogenation active metal comprises a Group VIII metal component, and Group VIII is selected from platinum and / or palladium. The hydrorefining catalyst is such as the FHUDS series catalyst developed by FRIPP. Components such as nitrides and aromatics in the material are removed.

[0025] Furthermore, the liquid phase hydrogenation reaction product is mixed with the light naphtha component 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) The present invention controls the reaction conditions of the first fixed bed reactor to react the small molecular sulfides in the coking liquefied gas with the olefins in the liquefied gas and / or the coking diesel, and converts the reaction products into high-boiling-point, large-molecule mercaptan sulfur, sulfide and other substances, while retaining most of the olefin components in the liquefied gas without undergoing saturation reaction, and using the coking diesel to directly take the generated high-boiling-point mercaptan sulfur and sulfide substances out of the reaction system, and then carry out the subsequent gas phase hydrogenation reactor to undergo hydrodesulfurization and olefin saturation reaction.

[0028] (2) The present invention can also utilize the relatively active diene components in coking diesel to react with mercaptans in liquefied gas for thioetherification reaction, and can remove dienes in oil products while pre-treating coking diesel, thereby avoiding the problem of dienes coking in heat exchangers, heating furnaces, reactors, etc. when coking diesel is processed alone. In addition, the introduction of coking diesel can provide more olefin substances for the thioetherification reaction, which is beneficial to the removal of mercaptan sulfur substances in the reactor. At the same time, the introduced coking diesel can also act as a stabilizer in the system, and can take out the macromolecular sulfide generated after the thioetherification reaction, which is beneficial to the forward progress of the thioetherification reaction.

[0029] (3) The present invention can be directly coupled with the coking unit fractionation tower. By reasonably utilizing the heat exchange of the side diesel fraction of the coking unit fractionation tower, the coking diesel fraction directly extracted can enter the first fixed bed reactor. The coking diesel feedstock can directly enter the hydrogenation reactor without passing through the tank area or buffer tank. This can greatly reduce the amount of diolefins and olefins in the feedstock that are generated as a result of the coking diesel being placed for too long, thereby delaying the increase in the bed pressure drop and the amount of carbon deposits in the catalyst in the device, thereby extending the operation cycle of the device.

[0030] (4) The present invention can remove the hydrogen sulfide generated in the second fixed bed reactor by arranging a separator between the second fixed bed reactor and the third fixed bed reactor, and can ensure that the material entering the third fixed bed reactor does not contain hydrogen sulfide, and avoid hydrogen sulfide from competing with aromatic hydrocarbons and nitrides for adsorption on the catalyst in the third fixed bed reactor, which is conducive to the hydrogenation dearomatization and denitrogenation reaction. At the same time, the arrangement of the separator can also separate the water present in the oil product and the water that may be generated by the reaction, ensuring that the water-free substances entering the third reactor can be avoided, and the strength of the catalyst can be prevented from being destroyed when the water content is high, affecting the activity and stability of the catalyst, and can also reduce the partial pressure of water in the gas phase, which is conducive to the hydrogenation reaction.

[0031] (5) The present invention combines two reactors, gas phase hydrogenation and liquid phase hydrogenation, and sets different reaction conditions according to the advantages of different reactors, so that the desulfurization reaction and the denitrogenation and dearomatization reaction react separately, thereby reducing the hydrogen consumption of the device. The second fixed bed reactor is a gas phase hydrogenation reactor. By loading a catalyst with high desulfurization activity, most of the sulfides in the raw materials can be removed under high temperature, low pressure, and a large hydrogen-to-oil ratio, and the olefins can be saturated. At the same time, due to the high temperature and low pressure, the denitrogenation and dearomatization reactions are inhibited, and the mutual inhibition problem of the desulfurization and denitrogenation reactions can be avoided, thereby improving the reaction efficiency. The third fixed bed reactor is a liquid phase hydrogenation reaction, which is conducive to the denitrogenation and dearomatization reactions under higher pressure conditions. At the same time, hydrogen is supplemented at the reactor inlet and in the middle of the bed to ensure the hydrogen required for the hydrogenation reaction of the device. A separator is set in the two reactors to remove small molecular substances and hydrogen sulfide substances in the material, which is conducive to the device to maintain the liquid phase. At the same time, the main macromolecular polycyclic aromatic hydrocarbons and macromolecular sulfides entering the third fixed bed reactor are macromolecular polycyclic aromatic hydrocarbons and macromolecular sulfides, thereby improving the conversion rate of the aromatic hydrogenation reaction. BRIEF DESCRIPTION OF THE DRAWINGS

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

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

[0034] The present invention is further described below in conjunction with the examples, but it should be understood that the protection scope of the present invention is not limited by the examples. In the present invention, unless otherwise clearly stated, percentages and percentage contents are all based on mass.

[0035] Combine the following Figure 1 The process flow of the present invention is described in detail.

[0036] The desulfurized coking liquefied gas component 1 is mixed with hydrogen 3 and fed from the lower part of the first fixed bed reactor 4, and the coking diesel 2 is fed from the upper part of the first fixed bed reactor 4. A thioetherification reaction occurs in the first fixed bed reactor 4. The gaseous product is extracted from the upper side line of the first fixed bed reactor 4, passes through a liquid separator 5, and a dry gas 6 and a liquefied gas 7 are separated. The liquid component flows out from the bottom of the first fixed bed reactor 4, passes through a heat exchanger 8, and enters a hydrogen mixer 9, is mixed with hydrogen 3, and enters a second fixed bed reactor 10 for a hydrogenation reaction. The reactor material enters a high fractionator 11, and the high fractionator enters a liquid separator 12 to analyze a high fraction gas 13 and light naphtha 14. The liquid component separated from the high fraction is pressurized by a booster pump 15, enters a hydrogen mixer 16, is mixed with hydrogen 3, and then enters a third fixed bed reactor 17 for a deep hydrogenation refining reaction. The diesel component 18 after the reaction enters a subsequent gas stripping fractionation device to obtain diesel and naphtha components.

[0037] Examples 1 to 3

[0038] The first fixed bed reactor is filled with 50mL of thioetherification catalyst C, the carrier of the catalyst is alumina-silicon dioxide, the active components are Mo and Co components, wherein the content of the carrier is 79%, the content of MoO3 is 16%, the content of CoO is 3%, and a small amount of auxiliary agent. The second fixed bed reactor is filled with 50mL of Mo-Co type catalyst A, the carrier of catalyst A is alumina, the active metal components are Mo and Co, the content of the carrier is 74%, the content of active metal Mo in terms of oxide is 23%, the content of active metal Co in terms of oxide is 5%, and the rest is the carrier. The third fixed bed reactor is filled with 50mL of Mo-Ni type catalyst B, the active metal components are Mo and Ni, and contain a small amount of auxiliary agent, the content of active metal Mo in terms of oxide is 24%, the content of active metal Ni in terms of oxide is 5%, and contain a small amount of auxiliary agent, and the rest is the carrier. The catalyst C is in the upper state of the first reactor, liquefied gas and hydrogen are fed from the lower part of the reactor, coking diesel is fed from the upper part of the first reactor, liquefied gas products are extracted from the upper side line of the first reactor, and the products are separated by a separator, a separator and a corresponding separator are arranged between the second fixed bed reactor and the third fixed bed reactor, and a stripping and distillation device is arranged after the third reactor.

[0039] Liquefied gas and coking diesel without hydrogen sulfide were used as raw materials. The properties of the catalyst are shown in Table 1, the properties of the coking diesel raw materials 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. 100 mL of Mo-Co type catalyst A was loaded in the reactor, and a separator, stripping and other processes were set after the reactor. Hydrogen was passed once. The diesel feedstock was the same as in Example 1-3, and the reaction process conditions and results are shown in Table 3.

[0042] Comparative Example 2

[0043] A conventional fixed bed hydrogenation reactor was used to react under normal conditions. The reactor was loaded with 60 mL of Mo-Ni type catalyst B and 40 mL of Mo-Co type catalyst A (same as in Example 1-3) from top to bottom. A separator, stripping and other processes were normally set after the reactor. Hydrogen was passed once. The diesel feedstock was the same as in Example 1-3, and 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 with 60 mL of Mo-Ni type catalyst B and 40 mL of Mo-Co type catalyst A (same as in Example 1-3) from top to bottom. High fractionation, low fractionation, stripping and other processes were normally set up after the reactor. The diesel raw material was the same as in Example 1-3, and the reaction process conditions and results were shown in Table 3.

[0046] Comparative Example 4

[0047] A conventional fixed bed reactor was used, and 50 mL of thioetherification catalyst C was loaded inside the reactor. The desulfurized coking liquefied gas and hydrogen entered the reactor from the top of the reactor.

[0048] Table 1 Physicochemical properties of catalysts

[0049] Catalyst No. Thioetherification 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 Shamrock Shamrock 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 the data <![CDATA[Density (20 °C), kg·m -3 > 856.8 Distillation range, ℃ 165～365 Sulfur content, μg / g 11640 Total aromatics, wt% 27.8 PAHs wt% 12.8

[0052] Table 3 Properties of liquefied gas raw materials

[0053] Liquefied gas properties Analyze the data Olefin content, v% 24.2 Total S, μg / g 168 Mercaptan 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 contents: (1) coking diesel extracted from the side line of the fractionation tower of the coking unit enters from the top of a first fixed bed reactor after heat exchange and cooling, and coking liquefied gas after hydrogen sulfide removal enters from the middle and lower part of the first fixed bed reactor, and coking diesel and coking liquefied gas undergo sulfide etherification reaction in a catalyst bed arranged in the middle and upper part of the reactor, and the gaseous material after the reaction flows out from the side line of the reactor top, 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 mixed with hydrogen and then enters a second fixed bed reactor for gas-phase hydrogenation desulfurization and deolefination reaction; (3) the effluent of the second fixed bed reactor is subjected to gas-liquid separation, the separated gaseous component is condensed to obtain a light naphtha component, the separated liquid component is pressurized and then enters a third fixed bed reactor for liquid-phase hydrogenation reaction, and the liquid-phase hydrogenation reaction product is stripped and fractionated to obtain a refined product.

2. The process according to claim 1, characterized in that: The coking liquefied gas after the removal of hydrogen sulfide has an olefin content of 5v% to 60v%, preferably an olefin content ≯50v%, and a mercaptan sulfur content of 20 to 500μg / g, preferably 50 to 200μg / g.

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

4. The process according to claim 1, characterized in that: The temperature of the coking diesel extracted from the side line of the distillation tower of the coking unit is 250~350℃. After heat exchange, it enters the first fixed bed reactor. The inlet temperature of the first 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 thioetherification catalyst; the operating conditions of the first fixed bed reactor are: reaction pressure 0.1~3.0MPa, preferably 0.52.0MPa, reaction temperature 50~300℃, preferably 80~180℃, liquid hourly volume space velocity 0.1~20h -1 , preferably 1~10h -1 .

6. The process according to claim 1, characterized in that: The catalyst loaded in the second fixed bed reactor includes a protective agent, a silicon scavenger and a hydrogenation refining catalyst; the gas phase hydrogenation reaction conditions in the second fixed bed reactor are as follows: the pressure is 0.1~3.0MPa, preferably 0.5~2.0MPa; the hydrogen-oil volume ratio is 20~600, preferably 50~200; the temperature is 250~400℃, preferably 300~380℃; the volume space velocity is 0.1~6.0, preferably 1.0~4.

0.

7. The process according to claim 1, characterized in that: The gaseous components separated in step (3) are condensed to obtain light naphtha components and hydrogen, ammonia and hydrogen sulfide, wherein the light naphtha components are used as ethylene cracking materials. After separation, the hydrogen sulfide generated by the reaction in the second reactor 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 liquid phase reaction conditions in the third fixed bed reactor are as follows: pressure 0.5~12.0MPa, preferably 4~10.0MPa; temperature 200~400°C, preferably 300~380°C; volume space velocity 0.1~4.0, preferably 0.5~2.

0.

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

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

Images