Hydrogenation process for coking gasoline and diesel oil

Abstract

The invention discloses a coking gasoline and diesel hydrogenation process which comprises the following steps: (1) pumping coking gasoline and diesel components from a side line of a fractionating tower of a coking device and hydrogen into a first reactor, and carrying out gas-phase desulfurization reaction; (2) the first reactor effluent is subjected to heat exchange and then enters a gas-liquid separator, a gas-phase component is upwards discharged out of the gas-liquid separator to obtain a hydrogenation light component, and a liquid-phase component is downwards discharged out of the gas-liquid separator; (3) liquid-phase components flowing out of the gas-liquid separator are pressurized and mixed with hydrogen and then enter a second reactor, and bottom effluents of the second reactor are hydrogenation heavy components; and mixing the hydrogenated heavy component with the hydrogenated light component, and feeding the mixture into a stripping tower to remove hydrogen sulfide to obtain refined naphtha and refined diesel oil products. According to the method, coking gasoline and diesel can be processed at the same time, the reaction severity is reduced by controlling the reaction conditions, chemical hydrogen consumption is reduced, and meanwhile energy consumption of the device is saved from the overall technological process.

Description

Technical Field

[0001] The invention belongs to the field of oil refining and chemical industry, and specifically relates to a coking gasoline and diesel hydrogenation process. Background Art

[0002] At present, the hydrorefining of diesel generally adopts a fixed bed hydrogenation process. With the increasingly stringent environmental protection regulations, in addition to requiring the sulfur content in refined diesel to be ≯10ppm, the National VI diesel quality standard proposed in 2019 requires polycyclic aromatic hydrocarbons ≯11%, and the National VI B standard implemented in 2023 requires polycyclic aromatic hydrocarbons ≯5%. However, from the perspective of the reaction mechanism of deep desulfurization and deep dearomatization, there are large differences in the requirements for the reaction environment between the two. For the desulfurization reaction, the removal of small molecular sulfur mainly follows the direct desulfurization route, that is, hydrogenolysis. Large molecular sulfides with low reaction activity mainly follow the hydrodesulfurization reaction route, that is, first hydrogenation of aromatic rings and then hydrogenolysis. In the upper part of the reactor, the hydrogenation of small molecular sulfides and nitrides and the conversion of some dicyclic aromatic hydrocarbons (large amounts of heat release) mainly occur. In the lower part of the reactor, the accumulation of heat and hydrogen sulfide makes the reaction environment high temperature and low hydrogen partial pressure. This reaction environment is not conducive to the hydrodesulfurization path of large molecular sulfur and the saturation of polycyclic aromatic hydrocarbons. Especially at the end of the reaction, the catalyst activity decays. Raising the temperature can make up for the loss of the catalyst's desulfurization activity, but it will further affect the dearomatization effect. In addition to the different requirements for the reaction environment, the competitive adsorption of aromatics on the catalyst surface also inhibits deep desulfurization, making it difficult for traditional hydrogenation technology to meet the dual requirements of ultra-deep desulfurization and efficient saturation of aromatics.

[0003] In addition, there is also the problem of large energy loss in the processing of coking oil products. Generally, after the side line components of coking gasoline and coking diesel are extracted from the delayed coking tower, the condensed liquid components enter the tank area or are directly piped to the downstream hydrogenation reactor. Before entering the reactor, they need to be heat exchanged and heated again to make the inlet temperature of the oil products meet the reaction requirements. At the same time, the condensation and heating process will cause energy waste and loss. Since the hydrogenation of coking naphtha is generally used as an ethylene feedstock and the hydrogenation of coking diesel is used to produce automotive diesel products, the hydrogenation depths of the two reaction systems are different. If they are processed together, there will be some excess in the properties of naphtha, resulting in waste of hydrogen. In addition, there is no restriction on the nitrogen content index of diesel products. In the conventional hydrogenation process, deep desulfurization will also lead to the preferential removal of nitrides, which also causes the problem of excessive reaction and waste of hydrogen.

[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 in a fractionation system. The invention has the following disadvantages: the reaction process 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. Summary of the invention

[0006] In view of the shortcomings of the prior art, the present invention provides a coking gasoline and diesel hydrogenation process. The process of the present invention can simultaneously process coking gasoline and diesel, reduce the severity of the reaction by controlling the reaction conditions, reduce chemical hydrogen consumption, and save energy consumption of the device from the overall process flow.

[0007] The coking gasoline and diesel hydrogenation process of the present invention comprises the following steps:

[0008] (1) The coking gasoline and diesel components are directly extracted from the side line of the coking unit fractionation tower as reaction raw materials, and enter the first reactor with hydrogen to carry out gas phase desulfurization reaction;

[0009] (2) The effluent from the first reactor enters a gas-liquid separator after heat exchange, wherein the gas phase component is discharged upward from the gas-liquid separator to obtain a hydrogenated light component, and the liquid phase component is discharged downward from the gas-liquid separator;

[0010] (3) The liquid phase component flowing out of the gas-liquid separator enters the second reactor after being pressurized and mixed with hydrogen, and the bottom effluent of the second reactor is the hydrogenated heavy component; the hydrogenated heavy component is mixed with the hydrogenated light component and enters the stripping tower to remove hydrogen sulfide, thereby obtaining refined naphtha and refined diesel products.

[0011] Furthermore, the properties of the coking gasoline and diesel extracted from the side line of the distillation tower of the coking device are as follows: the initial distillation point is 60-100°C, the final distillation point is 300-400°C; the S content is ≯20000μg / g, preferably ≯15000μg / g, the N content is ≯5000μg / g, preferably ≯3000μg / g, the polycyclic aromatic hydrocarbons content is ≯50wt%, and the olefin content is ≯50wt%. For example, the S content of the diesel can be exemplified but not limited to: 8000μg / g, 9000μg / g, 10000μg / g, 11500μg / g, etc., the N content can be exemplified but not limited to: 500μg / g, 1000μg / g, 1500μg / g, etc., the polycyclic aromatic hydrocarbon content can be exemplified but not limited to: 25wt%, 30wt%, 35wt%, 40wt%, 45wt%, etc., and the olefin content can be exemplified but not limited to: 20wt%, 30wt%, 35wt%, 40wt%, etc.

[0012] Furthermore, the first reactor and the second reactor are both fixed bed reactors.

[0013] Furthermore, the outlet temperature of the coking gasoline and diesel extracted from the side line is 280-380°C. In order to meet the reaction temperature requirement of the first reactor, it is generally necessary to heat exchange with the outlet logistics of the first reactor and then enter the first reactor. The operating conditions of the first reactor are as follows: pressure 0.1-2.8MPa, preferably 1.0-2.0MPa, temperature 260-400°C, preferably 300-380°C, hydrogen-to-oil volume ratio 50-500, preferably 50-300, volume space velocity 0.5-3.0h -1 , preferably 0.8 to 2.5 hours -1 .

[0014] Furthermore, the catalyst loaded in the first reactor can be a catalyst with desulfurization function, preferably a Mo-Co type hydrodesulfurization catalyst. Generally, an alumina-based carrier is used. Based on the mass of the catalyst, the content of Mo in terms of molybdenum oxide is 15wt% to 30wt%, and the content of Co in terms of cobalt oxide is 2wt% to 6wt%. The catalyst can also contain an auxiliary component, such as at least one of phosphorus, silicon, boron, magnesium, fluorine, etc., and the mass content in the catalyst is generally less than 6wt%. For example, the Mo-Co type diesel deep desulfurization catalyst developed by Sinopec Fushun Petrochemical Research Institute (FRIPP). For example, FHUDS-5, FHUDS-7, etc.

[0015] Furthermore, the effluent from the first reactor is cooled by a heat exchanger, which may be a conventional commercial heat exchanger, such as a tubular heat exchanger. The heat exchange can ensure that the heavy macromolecular substances (such as dicyclic or higher aromatic hydrocarbons, quinoline, indole nitrides, dibenzothiophene sulfides, etc.) in the gas phase material flowing out of the first reactor are liquefied, and the heat exchange temperature is generally 300-340°C.

[0016] Furthermore, the gas-liquid separator is used to establish a gas-liquid balance of the effluent after the heat exchanger, so that the gas phase component flows upward and the liquid phase component flows downward, wherein the proportion of the liquid phase effluent is 1m% to 40m%, preferably 10m% to 30m%.

[0017] Furthermore, the hydrogen mixing unit includes a booster pump and a hydrogen mixer, which are used to pressurize the liquid phase components and then supplement the hydrogen required for the reaction to meet the reaction conditions and hydrogen consumption of the second reactor.

[0018] Furthermore, the operating conditions of the second reactor are as follows: pressure 2.0-10.0 MPa, preferably 3.0-7.0 MPa, temperature 200-400°C, preferably 260-360°C, volume space velocity 0.1-5.0 h -1 , preferably 0.5~3.0h -1 .

[0019] Furthermore, the pressure of the second reactor is at least 1.0 MPa higher than the pressure of the first reactor, preferably 2.0 to 6.0 MPa higher. The temperature of the second reactor is at least 15°C lower than the temperature of the first reactor, preferably 5 to 40°C lower.

[0020] Furthermore, the catalyst loaded in the second reactor can be a catalyst with a hydrogenation and dearomatization function, such as a non-precious metal catalyst or a precious metal catalyst. The non-precious metal catalyst is a Mo-Ni type catalyst. Generally, an alumina-based carrier is used. Based on the mass of the catalyst, the content of Mo in terms of molybdenum oxide is 15wt% to 30wt%, and the content of Ni in terms of cobalt oxide is 2wt% to 5wt%. The catalyst can also contain an auxiliary component, such as at least one of phosphorus, silicon, boron, magnesium, fluorine, etc., and the mass content in the catalyst is generally less than 6wt%. The Mo-Ni type catalyst is such as FHUDS-10, FHUDS-6, FHUDS-8, etc. developed by Sinopec Fushun Petrochemical Research Institute (FRIPP); the precious metal catalyst is Pt, Pd, etc. as an active metal, such as FHDA-10 catalyst developed by Sinopec Fushun Petrochemical Research Institute (FRIPP).

[0021] Furthermore, the hydrogenated heavy components flowing out of the second reactor are mixed with the gas phase hydrogenated light components flowing out of the top of the gas-liquid separator, and enter the stripping system to remove hydrogen sulfide to obtain refined naphtha (ethylene feedstock) and refined diesel products.

[0022] Furthermore, the S content of the refined naphtha is less than 600 μg / g, the olefin content is less than 1.0%, the aromatic content in the diesel product is less than 5wt%, and the S content is less than 10 μg / g.

[0023] The present invention also provides a hydrogenation reaction system, comprising a coking device and a hydrogenation device, wherein the coking device comprises a coking tower and a fractionation tower, and the hydrogenation device comprises a first reactor, i.e., a gas phase reactor, a gas-liquid separator, a heat exchanger, a second reactor and a hydrogen mixing unit; and the lateral extraction pipeline of the fractionation tower is connected to the raw material inlet of the first reactor.

[0024] Compared with the prior art, the method of the present invention has the following advantages:

[0025] (1) The process of the present invention can be directly coupled with the fractionation tower of the coking unit, making full use of the heat of the material extracted from the side line to meet the requirements of the raw material entering the hydrogenation reactor, eliminating the need for a heating furnace, and saving energy and reducing consumption. The coking gasoline and diesel raw materials can directly enter the hydrogenation reactor without passing through the tank area or buffer tank, reducing the investment in the device.

[0026] (2) The process of the present invention, in terms of diesel hydrogenation, avoids the defect that deep desulfurization and dearomatization are placed in the same reaction system and the reaction conditions are difficult to be compatible. The inventors have found that by controlling the reaction conditions of the first reactor, only the hydrogenation removal reaction of sulfides in the diesel fraction can occur, thereby separating the desulfurization and dearomatization reactions, optimizing the reaction conditions respectively, and improving the efficiency of each reaction. In terms of naphtha hydrogenation, the selectivity of the reaction is improved. The removal of sulfides and olefins in the gas phase reactor avoids excessive hydrogenation of nitrides and reduces chemical hydrogen consumption.

[0027] (3) The process method of the present invention controls the desulfurization reaction to follow a reaction path with low hydrogen consumption by matching the use of a high hydrogenation desulfurization path catalyst in the first reactor, reduces the chemical hydrogen consumption of the desulfurization reaction, and can also reduce the hydrogen-to-oil ratio of the first reactor and reduce the severity of the reaction. At the same time, by controlling the reaction conditions of the first reactor, it is not only for gas phase desulfurization, but more importantly for better coordination with the second reactor, which is more conducive to subsequent deep dearomatization.

[0028] (4) The present invention uses a small amount of heat exchange under high temperature conditions for the effluent of the first reactor, retains a large amount of reaction heat under the condition of ensuring the liquefaction of the macromolecules, meets the feeding conditions of the second reactor, and saves energy. In addition, since only the liquefied macromolecules are further hydrogenated, the reaction space velocity is reduced and the reaction effect is improved.

[0029] (5) The hydrogen mixing unit of the present invention increases the pressure while replenishing the hydrogen consumed by the second reactor. The hydrogen flowing out of the gas phase reactor does not need to be recycled, thus eliminating the need for a circulating hydrogen compressor and significantly reducing the investment in device construction. BRIEF DESCRIPTION OF THE DRAWINGS

[0030] Figure 1 This is a schematic diagram of the coking gasoline and diesel hydrogenation process of the present invention;

[0031] Among them, coking gasoline, diesel raw materials and hydrogen-1; the first reactor-2, the first reactor effluent-3; the heat exchanger-4; the gas-liquid separator-5; the liquid phase component-6; the hydrogenated light component-7; the hydrogen mixing unit-8; the second reactor-9; the hydrogenated heavy component-10; the stripping and fractionation system-11; the refined naphtha product-12, the refined diesel product-13. DETAILED DESCRIPTION

[0032] The present invention will be further described below in conjunction with embodiments, but it should be understood that the protection scope of the present invention is not limited by the embodiments.

[0033] In the present invention, unless otherwise explicitly stated, percentages and contents are all based on mass.

[0034] Combine the following Figure 1 The process flow of the present invention is described in detail.

[0035] The coking gasoline and diesel raw materials and hydrogen 1 enter the first reactor 2, where a gas-phase desulfurization reaction occurs to obtain the first reactor effluent 3; the effluent enters the heat exchanger 4, where it is cooled down and then enters the gas-liquid separator 5, where the hydrogenated light component 7 is discharged upward from the gas-liquid separator 5, and the liquid component 6 is discharged downward from the gas-liquid separator 5, and then enters the second reactor 9 through the hydrogen mixing unit 8 to obtain a hydrogenated heavy component 10, and the hydrogenated light component 7 is mixed with the hydrogenated seed component 10 and enters the stripping tower 11 to finally obtain a refined diesel product 12 and a refined naphtha product 13.

[0036] Examples 1-3

[0037] Use Figure 1Schematic diagram of the process. Two 100mL fixed bed hydrogenation reactors are connected in series, namely the first reactor and the second reactor. A heat exchanger, a gas-liquid separator and a hydrogen mixing unit are set between the reactors. The first reactor is a gas phase hydrogenation reactor, loaded with 50mL Mo-Co type diesel hydrogenation catalyst A, and the second reactor is a liquid phase hydrogenation reactor, loaded with 50mL Mo-Ni type diesel hydrogenation catalyst B. A gas phase outlet is set above the gas-liquid separator, which is connected to the bottom effluent of the second reactor and enters the stripping tower together to obtain a refined diesel product.

[0038] Coking gasoline and diesel were used as raw materials. The properties of the catalyst are shown in Table 1, the properties of the raw oil are shown in Table 2, and the reaction process conditions and results are shown in Table 3.

[0039] Comparative Example 1

[0040] The conventional gasoline and diesel fixed bed hydrogenation process was adopted, and two hydrogenation reactors were set in series, namely reactor 1 and reactor 2, and a stripping tower was set between the two reactors to remove hydrogen sulfide. Reactor 1 was filled with 50mL Mo-Co type diesel hydrogenation catalyst A, and reactor 2 was filled with 50mL Mo-Ni type diesel hydrogenation catalyst B. After the reactor, high-fraction, low-fraction, stripping, fractionation and other processes were conventionally set to obtain diesel products and naphtha products. After removing hydrogen sulfide, hydrogen was pressurized and recycled by a circulating hydrogen compressor. The raw materials and catalyst properties are the same as those in Examples 1-3, and the reaction process conditions and results are shown in Table 3.

[0041] Comparative Example 2

[0042] Coking naphtha and coking diesel were hydrogenated separately. Reactor 1 is a coking naphtha hydrogenation reactor, and reactor 2 is a coking diesel hydrogenation reactor. A heating furnace is set in front of both reactors. Reactor 1 is filled with 50mL Mo-Co type diesel hydrogenation catalyst A, and reactor 2 is filled with 50mL Mo-Ni type diesel hydrogenation catalyst B. The two sets of devices share the stripping and fractionation system to obtain diesel products and naphtha products. The properties of the raw materials and catalysts are the same as those in Examples 1-3. The reaction process conditions and results are shown in Table 3

[0043] Comparative Example 3

[0044] The same process flow as in Examples 1 to 3 was adopted, except that the hydrogen-to-oil ratio at the inlet of the gas phase reactor was increased, the hydrogen mixing unit before the liquid phase reactor was cancelled, and only a liquid phase booster pump was used instead. Reactor 1 was filled with 50 mL of Mo-Co type diesel hydrogenation catalyst A, and reactor 2 was filled with 50 mL of Mo-Ni type diesel hydrogenation catalyst B. The properties of the raw materials and catalysts were the same as in Examples 1-3, and the reaction process conditions and results are shown in Table 3.

[0045] Table 1 Physicochemical properties of catalysts

[0046] Catalyst No. A B Active Metals Mo-Co Mo-Ni <![CDATA[MoO3,wt%]]> 20 24 NiO or CoO, wt% 3.5 5.0 shape Shamrock Shamrock Diameter, mm 1.2 1.2 <![CDATA[Specific surface area, m 2 ·g -1 > 180 180 <![CDATA[Pore volume, mL·g -1 > 0.35 0.35

[0047] Table 2 Raw oil properties

[0048] Oil properties <![CDATA[Density (20 °C), g·cm -3 > 0.84 Distillation range, ℃ 60～360 <![CDATA[S,μg·g -1 ]]> 10350 <![CDATA[N,μg·g -1 ]]> 2725 Aromatics, wt% 43.2 PAHs, wt% 18.7 Monocyclic aromatic hydrocarbons, wt% 24.5 Olefins, wt% 26

[0049] Table 3 Hydrogenation process conditions and results

[0050]

[0051]

[0052] As can be seen from Table 3, Comparative Example 1 adopts conventional fixed bed hydrogenation technology and traditional catalyst loading system, and the pressure of Reactor 1 is relatively high, resulting in simultaneous desulfurization, denitrogenation and dearomatization, which affect each other. The reaction conditions of the two reactors cannot be optimized for a certain reaction, resulting in poor reaction effect and high hydrogen consumption, hydrogen recycling, and high energy consumption. However, under the premise that the pressure of the first reactor is significantly reduced, the reaction system of the present invention can optimize the reaction conditions in a targeted manner to obtain a better dearomatization effect because the desulfurization and dearomatization and denitrogenation are divided into two reaction systems, and the nitrides that do not need hydrogenation are better retained in both naphtha and diesel fractions, avoiding the restriction of nitrides on hydrogenation reactions and improving reaction efficiency; sulfides are removed through the hydrogenolysis path, and nitrides are hydrogenated in small amounts to reduce hydrogen consumption; hydrogen does not need to be recycled, reducing energy consumption.

[0053] Comparative Example 2: Hydrogenation of coker naphtha and coker diesel separately can reduce the severity of the reaction of coker naphtha, but the heat of the oil extracted from the side line of the coking unit is still not fully utilized, and the energy consumption is high. In addition, a large amount of nitrides are still removed during the hydrogenation of naphtha and diesel, resulting in high hydrogen consumption.

[0054] In Comparative Example 3, no hydrogen mixing unit is provided in front of the liquid phase reactor, and sufficient hydrogen source is provided by increasing the hydrogen-to-oil ratio at the inlet of the gas phase reactor. However, the high hydrogen-to-oil ratio at the inlet of the gas phase reactor will lead to low oil phase partial pressure in the separator and insufficient liquefaction, that is, it is difficult for all polycyclic aromatic hydrocarbons to enter the liquid phase components for liquid phase hydrogenation reaction, resulting in a high content of polycyclic aromatic hydrocarbons in the refined oil. In addition, a compressor is required for hydrogen circulation, and the energy consumption is still higher than that of the embodiment.

Claims

1. A coking gasoline and diesel hydrogenation process, characterized in that The method comprises the following contents: (1) extracting the coking gasoline and diesel components from the side line of the fractionation tower of the coking unit directly as reaction raw materials, and entering the first reactor with hydrogen to carry out gas phase desulfurization reaction; (2) the effluent of the first reactor enters the gas-liquid separator after heat exchange, wherein the gas phase components are discharged upward from the gas-liquid separator to obtain hydrogenated light components, and the liquid phase components are discharged downward from the gas-liquid separator; (3) the liquid phase components flowing out of the gas-liquid separator are pressurized and mixed with hydrogen and enter the second reactor, and the bottom effluent of the second reactor is the hydrogenated heavy components; the hydrogenated heavy components are mixed with the hydrogenated light components and enter the stripping tower to remove hydrogen sulfide to obtain refined naphtha and refined diesel products.

2. The hydrogenation process according to claim 1, characterized in that: The properties of the coking gasoline and diesel extracted from the side line of the distillation tower of the coking device are as follows: the initial distillation point is 60-100°C, the final distillation point is 300-400°C; the S content is ≯20000μg / g, preferably ≯15000μg / g, the N content is ≯5000μg / g, preferably ≯3000μg / g, the polycyclic aromatic hydrocarbon content is ≯50wt%, and the olefin content is ≯50wt%.

3. The hydrogenation process according to claim 1, characterized in that: The outlet temperature of the coking gasoline and diesel extracted from the side line is 280-380°C. In order to meet the reaction temperature requirement of the first reactor, it needs to be heated with the outlet logistics of the first reactor before entering the first reactor; the operating conditions of the first reactor are as follows: pressure 0.1-2.8MPa, preferably 1.0-2.0MPa, temperature 260-400°C, preferably 300-380°C, hydrogen-to-oil volume ratio 50-500, preferably 50-300, volume space velocity 0.5-3.0h -1 , preferably 0.8~2.5h -1 .

4. The hydrogenation process according to claim 1, characterized in that: The catalyst loaded in the first reactor uses a catalyst with desulfurization function; preferably a Mo-Co type hydrodesulfurization catalyst, using an alumina-based carrier, based on the mass of the catalyst, the content of Mo calculated as molybdenum oxide is 15wt%~30wt%, and the content of Co calculated as cobalt oxide is 2wt%~6wt%.

5. The hydrogenation process according to claim 1, characterized in that: The effluent from the first reactor is cooled to 300-340° C. by a heat exchanger.

6. The hydrogenation process according to claim 1, characterized in that: The gas-liquid separator is used to establish a gas-liquid balance of the effluent after the heat exchanger, so that the gas phase components flow upward and the liquid phase components flow downward; The proportion of liquid effluent is 1m%~40m%, preferably 10m%~30m%.

7. The hydrogenation process according to claim 1, characterized in that: The hydrogen mixing unit includes a booster pump and a hydrogen mixer, which are used to supplement the hydrogen required for the reaction on the basis of pressurizing the liquid phase components to meet the reaction conditions and hydrogen consumption of the second reactor.

8. The hydrogenation process according to claim 1, characterized in that: The operating conditions of the second reactor are as follows: pressure 2.0-10.0 MPa, preferably 3.0-7.0 MPa, temperature 200-400°C, preferably 260-360°C, volume space velocity 0.1-5.0 h -1 , preferably 0.5~3.0h -1 .

9. The hydrogenation process according to claim 1, characterized in that: The pressure of the second reactor is at least 1.0 MPa higher than that of the first reactor, preferably 2.0-6.0 MPa higher; the temperature of the second reactor is at least 15° C. lower than that of the first reactor, preferably 5-40° C. lower.

10. The hydrogenation process according to claim 1, characterized in that: The catalyst loaded in the second reactor is a catalyst having a hydrogenation and dearomatization function.

11. The hydrogenation process according to claim 1, characterized in that: The refined naphtha has an S content of less than 600 μg / g and an olefin content of less than 1.0%. The refined diesel product has an aromatic content of less than 5 wt% and an S content of less than 10 μg / g.

12. A hydrogenation reaction system for the hydrogenation process according to claim 1, characterized in that: The invention comprises a coking device and a hydrogenation device, wherein the coking device comprises a coking tower and a fractionation tower, and the hydrogenation device comprises a first reactor, i.e. a gas phase reactor, a gas-liquid separator, a heat exchanger, a second reactor and a hydrogen mixing unit; and a lateral extraction pipeline of the fractionation tower is connected to a raw material inlet of the first reactor.

Images