CN119081732A

CN119081732A - A coking gasoline and diesel hydrogenation process

Info

Publication number: CN119081732A
Application number: CN202310655363.8A
Authority: CN
Inventors: 代萌; 徐大海; 丁贺; 牛世坤; 李扬
Original assignee: Sinopec Dalian Petrochemical Research Institute Co ltd; China Petroleum and Chemical Corp
Current assignee: Sinopec Dalian Petrochemical Research Institute Co ltd; China Petroleum and Chemical Corp
Priority date: 2023-06-05
Filing date: 2023-06-05
Publication date: 2024-12-06

Abstract

The invention discloses a coking gasoline and diesel hydrogenation process which comprises the following steps of (1) extracting coking gasoline and diesel components and hydrogen from a side line of a fractionating tower of a coking device, enabling the coking gasoline and diesel components and hydrogen to enter a first reactor for gas phase desulfurization reaction, (2) enabling a first reactor effluent to enter a gas-liquid separator after heat exchange, enabling the gas phase components to be discharged upwards from the gas-liquid separator to obtain hydrogenation light components, enabling the liquid phase components to be discharged downwards from the gas-liquid separator, and (3) enabling the liquid phase components flowing out of the gas-liquid separator to enter a second reactor after pressurization and hydrogen mixing, enabling the effluent at the bottom of the second reactor to be hydrogenation heavy components, enabling the hydrogenation heavy components to be mixed with the hydrogenation light components, enabling the hydrogenation heavy components to enter a stripping tower to remove hydrogen sulfide, and obtaining refined naphtha and refined diesel products. The method can process coked gasoline and diesel oil simultaneously, reduce reaction severity by controlling reaction conditions, reduce chemical hydrogen consumption, and save energy consumption of the device in the overall process flow.

Description

Coking gasoline and diesel oil hydrogenation process

Technical Field

The invention belongs to the field of oil refining and chemical industry, and particularly relates to a coking gasoline and diesel hydrogenation process.

Background

At present, a fixed bed hydrogenation process is generally adopted for the hydrofining of diesel oil. With the increasing strictness of environmental regulations, besides the requirement that the sulfur content in refined diesel is no more than 10ppm, in the national VI diesel quality standard proposed in 2019, no more than 11% of polycyclic aromatic hydrocarbon is required, and in the national VIB standard implemented in 2023, no more than 5% of polycyclic aromatic hydrocarbon is required. However, there is a great difference in the requirements of the reaction environment from the reaction mechanism of deep desulfurization and deep dearomatization. For desulfurization reactions, the removal of small molecule sulfur mainly follows the direct desulfurization route, i.e., hydrogenolysis desulfurization. The macromolecular sulfides with lower reactivity mainly follow the hydrodesulfurization reaction route, namely aromatic ring hydrogenation is carried out first, and then hydrogenolysis desulfurization is carried out. In the upper part of the reactor, the hydrogenation of small molecule sulfides, nitrides and the conversion of part of the bicyclic aromatic hydrocarbons (a large exotherm) mainly take place. The accumulation of heat and hydrogen sulfide at the lower part of the reactor ensures that the reaction environment is high temperature and low in hydrogen partial pressure, and the reaction environment is unfavorable for the hydrodesulfurization path of macromolecular sulfur and the saturation of polycyclic aromatic hydrocarbon. Especially, when the reaction is finished, the activity of the catalyst is attenuated, and the temperature is raised to compensate the desulfurization activity loss of the catalyst, but the dearomatization effect is further influenced. Besides different requirements of the two on reaction environment, competitive adsorption of aromatic hydrocarbon on the surface of the catalyst also has an inhibition effect on deep desulfurization, so that the traditional hydrogenation technology has difficulty in meeting the double requirements of ultra deep desulfurization and efficient saturation of aromatic hydrocarbon.

In addition, the processing of the coked oil product has the problem of larger energy loss. Generally, coker gasoline and coker diesel oil are pumped out of a delayed coker side line component, and condensed liquid phase components enter a tank area or directly are conveyed to a downstream hydrogenation reactor through a pipe, and heat exchange and heating are needed again before the condensed liquid phase components enter the reactor, so that the inlet temperature of the oil product reaches the reaction requirement, and meanwhile, the energy consumption is wasted and lost in the condensation and heating process. Because coked naphtha hydrogenation is generally used as an ethylene raw material, coked diesel hydrogenation is used for producing automotive diesel products, the hydrogenation depths of two reaction systems are different, and if the two reaction systems are processed together, the property of naphtha is excessive, so that hydrogen consumption is wasted. In addition, the diesel oil product has no limit requirement on the nitrogen content index, and in the conventional hydrogenation process, the advanced desulfurization can also lead to the preferential removal of nitride, and the problem of hydrogen waste caused by excessive reaction is also caused.

CN106190234a discloses a hydrodesulfurization refining process for coked diesel, which adopts a fixed bed reactor, wherein the reactor is filled with a catalyst with hydrodesulfurization and denitrification activity, and the active components of the catalyst are nitriding and/or carbonized metals, wherein the metal components are molybdenum, tungsten and the like. The sulfide in the raw material can be controlled below 5ppm under the conditions of 300-450 ℃, hydrogen partial pressure of 6.0-9.0 MPa, volume space velocity of 1.5-3 h ^-1 and hydrogen oil volume ratio of 500. However, the process flow is complex, raw materials need to enter a raw material buffer tank of a reaction unit for reaction after passing through a tank area and prefractionation, and the equipment investment is large.

CN101003745A discloses a method for producing high quality diesel oil from coked full distillate, the coked full distillate is separated into light fraction and heavy fraction, hydrofining and hydro-upgrading are respectively carried out, and the reaction products of the two are sent into a fractionation system together to be separated into naphtha and diesel oil. The invention firstly has complex reaction flow, and secondly, the heavy distillate oil enters a hydro-upgrading reaction zone, and the nitride in the heavy distillate oil can seriously influence the activity of the modifier.

Disclosure of Invention

Aiming at the defects of the prior art, the invention provides a coking gasoline and diesel hydrogenation process. The process can process coked gasoline and diesel oil simultaneously, reduce reaction severity by controlling reaction conditions, reduce chemical hydrogen consumption, and save energy consumption of the device in the overall process flow.

The invention relates to a coking gasoline and diesel hydrogenation process, which comprises the following steps:

(1) The coking gasoline and diesel components are directly taken as reaction raw materials and enter a first reactor together with hydrogen to carry out gas-phase desulfurization reaction;

(2) The effluent of the first reactor enters a gas-liquid separator after heat exchange, wherein the gas phase component is discharged upwards from the gas-liquid separator to obtain a hydrogenation light component, and the liquid phase component is discharged downwards from the gas-liquid separator;

(3) The liquid phase component flowing out of the gas-liquid separator enters a second reactor after being pressurized and hydrogen mixed, the effluent at the bottom of the second reactor is a heavy hydrogenation component, the heavy hydrogenation component is mixed with the light hydrogenation component, and the mixture enters a stripping tower to remove hydrogen sulfide, so that refined naphtha and refined diesel oil products are obtained.

Further, the property of the coker fractionating tower side-drawn coker gasoline and diesel oil is that the initial distillation point is 60-100 ℃ and the final distillation point is 300-400 ℃, the S content is no more than 20000 mug/g, preferably no more than 15000 mug/g, the N content is no more than 5000 mug/g, preferably no more than 3000 mug/g, the polycyclic aromatic hydrocarbon content is no more than 50wt%, and the olefin content is no more than 50wt%. For example, the S content of the diesel fuel may be, but not limited to, 8000. Mu.g/g, 9000. Mu.g/g, 10000. Mu.g/g, 11500. Mu.g/g, etc., the N content may be, but not limited to, 500. Mu.g/g, 1000. Mu.g/g, 1500. Mu.g/g, etc., the polycyclic aromatic hydrocarbon content may be, but not limited to, 25wt%, 30wt%, 35wt%, 40wt%, 45wt%, etc., and the olefin content may be, but not limited to, 20wt%, 30wt%, 35wt%, 40wt%, etc.

Further, the first reactor and the second reactor are both fixed bed reactors.

Further, the outlet temperature of the side-drawn coking gasoline and diesel oil is 280-380 ℃, and in order to meet the reaction temperature requirement of the first reactor, the coking gasoline and diesel oil generally needs to enter the first reactor after heat exchange and temperature rising with the outlet material flow of the first reactor, and the operation condition of the first reactor is that the pressure is 0.1-2.8 MPa, preferably 1.0-2.0 MPa, the temperature is 260-400 ℃, preferably 300-380 ℃, the hydrogen-oil volume ratio is 50-500, preferably 50-300, and the volume space velocity is 0.5-3.0 h ^-1, preferably 0.8-2.5 h ^-1.

Further, the catalyst filled in the first reactor may be a catalyst with a desulfurization function, preferably a Mo-Co type hydrodesulfurization catalyst, generally an alumina-based carrier is used, the content of Mo is 15wt% to 30wt% calculated as molybdenum oxide based on the mass of the catalyst, and the content of Co is 2wt% to 6wt% calculated as cobalt oxide. The catalyst may further 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 6wt% or less. Such as Mo-Co type diesel oil deep desulfurization catalyst developed by China petrochemical smoothing petrochemical institute (FRIPP). For example FHUDS-5, FHUDS-7, etc.

Further, the first reactor effluent is cooled via a heat exchanger, which may be a conventional commercial heat exchanger, such as a tubular heat exchanger. The heat exchange can ensure that macromolecular heavy substances (such as aromatic hydrocarbon with more than double rings, quinoline, indole nitrides, dibenzothiophene sulfides and the like) in the gas-phase material flowing out of the first reactor are liquefied, and the heat exchange is generally carried out to 300-340 ℃.

Further, the gas-liquid separator is used for establishing the gas-liquid balance of the effluent after the heat exchanger, so that the upward outflow of the gas phase components and the downward outflow of the liquid phase components are realized. Wherein the proportion of the liquid phase effluent is 1-40 m%, preferably 10-30 m%.

Further, the hydrogen mixing unit comprises a booster pump and a hydrogen mixer, and is used for boosting the liquid phase component and supplementing hydrogen required by the reaction, so that the reaction conditions and hydrogen consumption of the second reactor are met.

Further, the second reactor is operated under the conditions that the pressure is 2.0-10.0 MPa, preferably 3.0-7.0 MPa, the temperature is 200-400 ℃, preferably 260-360 ℃, and the volume space velocity is 0.1-5.0 h ^-1, preferably 0.5-3.0 h ^-1.

Further, the pressure of the second reactor is at least 1.0MPa higher than that of the first reactor, preferably 2.0-6.0 MPa higher. The temperature of the second reactor is at least 15 ℃ lower than the temperature of the first reactor, preferably 5-40 ℃ lower.

Further, the catalyst filled in the second reactor may be a catalyst having a hydrodearomatics function, such as a non-noble metal catalyst or a noble metal catalyst, where the non-noble metal catalyst is a Mo-Ni type catalyst, typically, an alumina-based carrier is used, based on the mass of the catalyst, mo is 15wt% to 30wt% based on molybdenum oxide, and Ni is 2wt% to 5wt% based on cobalt oxide. The catalyst may further 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 6wt% or less. The Mo-Ni catalyst is FHUDS-10, FHUDS-6, FHUDS-8 and the like developed by China petrochemical and smooth petrochemical institute (FRIPP), and the noble metal catalyst is FHDA-10 developed by China petrochemical and smooth petrochemical institute (FRIPP) and takes Pt, pd and the like as active metals.

Further, the heavy hydrogenation component flowing out of the second reactor is mixed with the gas phase light hydrogenation component flowing out of the top of the gas-liquid separator, and enters a stripping system to remove hydrogen sulfide, thus obtaining refined naphtha (ethylene raw material) and refined diesel oil products.

Further, the refined naphtha has an S content of 600 μg/g or less, an olefin content of 1.0% or less, an aromatic hydrocarbon content of 5wt% or less, and an S content of 10 μg/g or less in a diesel product.

The invention also provides a hydrogenation reaction system which comprises a coking device and a hydrogenation device, wherein the coking device comprises a coking tower and a fractionating tower, the hydrogenation device comprises a first reactor, namely 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 fractionating tower is connected with a raw material inlet of the first reactor.

Compared with the prior art, the method has the following advantages:

(1) The process method can be directly coupled with the coking unit fractionating tower, fully utilizes the heat of the side-draw materials, meets the requirement that the raw materials enter the inlet of the hydrogenation reactor, omits a heating furnace, saves energy and reduces consumption. The coked gasoline and diesel oil raw materials can directly enter the hydrogenation reactor without passing through a tank area or a buffer tank, so that the investment of the device is reduced.

(2) The process method of the invention avoids the defects of deep desulfurization and dearomatization in the same reaction system and difficult compatibility of reaction conditions in the aspect of diesel hydrogenation. The inventor researches and discovers that only hydrogenation removal reaction of sulfides in diesel oil fraction can be carried out by controlling the reaction conditions of the first reactor, thereby realizing separation of desulfurization and dearomatization reaction, respectively optimizing the reaction conditions and improving the respective reaction efficiency. In the aspect of naphtha hydrogenation, the selectivity of the reaction is improved. Sulfide and olefin are removed in the gas phase reactor, so that excessive hydrogenation of nitride is avoided, and chemical hydrogen consumption is reduced.

(3) According to the process method, the high hydrogenolysis desulfurization path catalyst is matched in the first reactor, so that the desulfurization reaction is controlled to more follow the reaction path with low hydrogen consumption, the chemical hydrogen consumption of the desulfurization reaction is reduced, the hydrogen-oil ratio of the first reactor can be reduced, and the reaction severity is reduced. Meanwhile, by controlling the reaction conditions of the first reactor, the method is not only used for gas-phase desulfurization, but also used for better matching with the second reactor, thereby being more beneficial to subsequent deep dearomatization.

(4) The invention realizes meeting the feeding condition of the second reactor and saving energy consumption by adopting a small amount of heat exchange under the high temperature condition for the effluent of the first reactor and largely retaining the reaction heat under the condition of ensuring that macromolecules are liquefied. And as only liquefied macromolecules are further hydrogenated, the reaction space velocity is reduced, and the reaction effect is improved.

(5) The hydrogen mixing unit of the invention supplements hydrogen consumption of the second reactor while pressurizing, hydrogen flowing out of the gas phase reactor can be used without circulation, a circulating hydrogen compressor is omitted, and the construction investment of the device is greatly reduced.

Drawings

FIG. 1 is a schematic diagram of a coking gasoline and diesel hydrogenation process flow of the present invention;

The method comprises the steps of coking a gasoline and diesel raw material and hydrogen gas-1, a first reactor-2, a first reactor effluent-3, a heat exchanger-4, a gas-liquid separator-5, a liquid phase component-6, a hydrogenation light component-7, a hydrogen mixing unit-8, a second reactor-9, a hydrogenation heavy component-10, a stripping and fractionating system-11, a refined naphtha product-12 and a refined diesel product-13.

Detailed Description

The present invention will be further described with reference to examples, but it should be understood that the scope of the present invention is not limited by the examples.

In the present invention, percentages and percentages are by mass unless explicitly stated otherwise.

The process flow of the present invention is described in detail below in conjunction with fig. 1.

The coking gasoline and diesel oil raw materials and hydrogen 1 enter a first reactor 2 to generate gas phase desulfurization reaction to obtain a first reactor effluent 3, enter a heat exchanger 4, enter a gas-liquid separator 5 after being cooled by the heat exchanger 4, the hydrogenated light component 7 is discharged upwards from the gas-liquid separator 5, the liquid component 6 is discharged downwards from the gas-liquid separator 5, then enter a second reactor 9 through a hydrogen mixing unit 8 to obtain a hydrogenated heavy component 10, the hydrogenated light component 7 and the hydrogenated light component 10 are mixed and enter a stripping tower 11, and finally a refined diesel oil product 12 and a refined naphtha product 13 are obtained.

Examples 1 to 3

A schematic flow chart as in fig. 1 is employed. Two 100mL fixed bed hydrogenation reactors are connected in series, namely a first reactor and a second reactor. The heat exchanger, the gas-liquid separator and the hydrogen mixing unit are arranged among the reactors. The first reactor is a gas-phase hydrogenation reactor filled with 50mLMo-Co type diesel hydrogenation catalyst A, and the second reactor is a liquid-phase hydrogenation reactor filled with 50mLMo-Ni type diesel hydrogenation catalyst B. And a gas phase outlet is arranged above the gas-liquid separator, and the gas phase outlet and the effluent at the bottom of the second reactor are connected and then enter a stripping tower together to obtain a refined diesel product.

Coked gasoline and diesel oil are used as raw materials, the properties of the catalyst are shown in table 1, the properties of the raw material oil are shown in table 2, and the reaction process conditions and results are shown in table 3.

Comparative example 1

The conventional gasoline and diesel fixed bed hydrogenation process flow is adopted, two hydrogenation reactors are connected in series, namely a reactor 1 and a reactor 2, and a stripping tower is arranged between the two reactors to remove hydrogen sulfide. Reactor 1 was charged with 50mLMo-Co type diesel hydrogenation catalyst A and reactor 2 was charged with 50mLMo-Ni type diesel hydrogenation catalyst B. The reactor is provided with high-fraction, low-fraction, steam stripping, fractionation and other processes to obtain diesel oil product and naphtha product. The hydrogen is pressurized and recycled through a recycle hydrogen compressor after hydrogen sulfide is removed. The raw materials and the catalyst properties were the same as in examples 1-3, and the reaction conditions and results are shown in Table 3.

Comparative example 2

The coking naphtha and the coking diesel oil are adopted for hydrogenation respectively. The reactor 1 is a coking naphtha hydrogenation reactor, the reactor 2 is a coking diesel hydrogenation reactor, and heating furnaces are arranged in front of the two reactors. Reactor 1 was charged with 50mLMo-Co type diesel hydrogenation catalyst A and reactor 2 was charged with 50mLMo-Ni type diesel hydrogenation catalyst B. The two sets share a stripping and fractionating system to obtain a diesel product and a naphtha product. The raw materials and the catalyst properties are the same as those of examples 1 to 3, and the reaction conditions and the results are shown in Table 3

Comparative example 3

The same process flow as in examples 1-3 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 eliminated, and only the liquid phase booster pump was used instead. Reactor 1 was charged with 50mLMo-Co type diesel hydrogenation catalyst A and reactor 2 was charged with 50mLMo-Ni type diesel hydrogenation catalyst B. The raw materials and the catalyst properties were the same as in examples 1-3, and the reaction conditions and results are shown in Table 3.

TABLE 1 catalyst physicochemical Properties

Catalyst numbering	A	B
			Reactive metal	Mo-Co	Mo-Ni
MoO₃,wt%	20	24
			NiO or CoO, wt%	3.5	5.0
Shape and shape	Clover with three leaves	Clover with three leaves
			Diameter of mm	1.2	1.2
Specific surface area, m ²·g^-1	180	180
			Pore volume, mL.g ^-1	0.35	0.35

TABLE 2 oil Properties of raw materials

Oil Properties
		Density (20 ℃), g.cm ^-3	0.84
The distillation range, C	60~360
		S,μg·g^-1	10350
N,μg·g^-1	2725
		Aromatic hydrocarbon, wt%	43.2
Polycyclic aromatic hydrocarbon, wt%	18.7
		Monocyclic aromatic hydrocarbon, wt%	24.5
Olefins, wt%	26

TABLE 3 hydrogenation process conditions and results

As can be seen from Table 3, in comparative example 1, the conventional fixed bed hydrogenation technology and the conventional catalyst loading system are adopted, the pressure of the reactor 1 is higher, the desulfurization, denitrification and dearomatization are simultaneously carried out and are mutually influenced, the reaction conditions of the two reactors cannot be optimized for a certain reaction, the reaction effect is poor, the hydrogen consumption is higher, the hydrogen is recycled, and the energy consumption is high. On the premise that the pressure of the first reactor is obviously reduced, the reaction system divides desulfurization and dearomatization and denitrification into two reaction systems, so that the reaction conditions can be pertinently optimized, a better dearomatization effect is obtained, hydrogenated nitrides are better reserved in naphtha and diesel fractions, the limitation of the nitrides on hydrogenation reaction is avoided, the reaction efficiency is improved, sulfides are removed through a hydrogenolysis path, a small amount of nitrides are hydrogenated, the hydrogen consumption is reduced, the hydrogen is not required to be recycled, and the energy consumption is reduced.

In comparative example 2, the coker naphtha and the coker diesel are hydrogenated respectively, so that the reaction severity of the coker naphtha can be reduced, the heat of the side oil extracted from the coking device is not fully utilized until then, and the energy consumption is high. And a large amount of nitride is still removed in the hydrogenation process of naphtha and diesel, so that the hydrogen consumption is high.

Comparative example 3a hydrogen mixing unit was not provided before the liquid phase reactor, and a sufficient hydrogen source was provided by increasing the hydrogen-to-oil ratio at the inlet of the gas phase reactor. However, the high hydrogen-oil ratio at the inlet of the gas phase reactor can lead to low oil phase partial pressure in the separator and insufficient liquefaction, namely, polycyclic aromatic hydrocarbon is difficult to fully enter liquid phase components for liquid phase hydrogenation reaction, so that the content of polycyclic aromatic hydrocarbon in refined oil is higher. In addition, hydrogen recycling requires the use of a compressor, and energy consumption is still higher than in the examples.

Claims

1. A coking gasoline and diesel hydrogenation process is characterized by comprising the following steps of (1) directly taking coking gasoline and diesel components as reaction raw materials at a side line of a fractionating tower of a coking device, enabling the coking gasoline and diesel components and hydrogen to enter a first reactor for gas phase desulfurization reaction, (2) enabling effluent of the first reactor to enter a gas-liquid separator after heat exchange, wherein the gas phase components are upwards discharged from the gas-liquid separator to obtain hydrogenation light components, and enabling the liquid phase components discharged from the gas-liquid separator to downwards discharge from the gas-liquid separator, (3) enabling the liquid phase components flowing out of the gas-liquid separator to enter a second reactor after pressurization and hydrogen mixing, enabling the effluent at the bottom of the second reactor to be hydrogenation heavy components, mixing the hydrogenation heavy components with the hydrogenation light components, and enabling the hydrogenation heavy components to enter a stripping tower to remove hydrogen sulfide to obtain refined naphtha and refined diesel products.

2. The hydrogenation process according to claim 1, wherein the coker fractionator side draw coker gas oil has the characteristics of initial distillation point of 60-100 ℃ and final distillation point of 300-400 ℃, S content of 20000 μg/g, preferably no more than 15000 μg/g, N content of 5000 μg/g, preferably no more than 3000 μg/g, polycyclic aromatic hydrocarbon content of no more than 50wt% and olefin content of no more than 50wt%.

3. The hydrogenation process according to claim 1, wherein the outlet temperature of the side-drawn coker gasoline and diesel oil is 280-380 ℃, and the operating conditions of the first reactor are as follows, wherein the pressure is 0.1-2.8 MPa, preferably 1.0-2.0 MPa, the temperature is 260-400 ℃, preferably 300-380 ℃, the hydrogen-oil volume ratio is 50-500, preferably 50-300, and the volume space velocity is 0.5-3.0 h ^-1, preferably 0.8-2.5 h ^-1.

4. The hydrogenation process according to claim 1, wherein the catalyst packed in the first reactor is a catalyst having a desulfurization function, preferably a Mo-Co hydrodesulfurization catalyst, and the catalyst comprises 15 to 30wt% of Mo and 2 to 6wt% of Co, based on the mass of the catalyst.

5. The hydrogenation process of claim 1, wherein the effluent from the first reactor is cooled to 300-340 ℃ by a heat exchanger.

6. The hydrogenation process according to claim 1, wherein the gas-liquid separator is used for establishing the gas-liquid balance of the effluent after the heat exchanger to realize the upward outflow of the gas phase component and the downward outflow of the liquid phase component, wherein the ratio of the liquid phase effluent is 1m% -40 m%, preferably 10 m% -30 m%.

7. The hydrogenation process according to claim 1, wherein the hydrogen mixing unit comprises a booster pump and a hydrogen mixer for supplementing hydrogen required by the reaction on the basis of boosting the liquid phase component, thereby meeting the reaction conditions and hydrogen consumption of the second reactor.

8. The hydrogenation process according to claim 1, wherein the second reactor is operated at a pressure of 2.0-10.0 MPa, preferably 3.0-7.0 MPa, a temperature of 200-400 ℃, preferably 260-360 ℃, and a volume space velocity of 0.1-5.0 h ^-1, preferably 0.5-3.0 h ^-1.

9. The hydrogenation process according to claim 1, wherein the pressure in the second reactor is at least 1.0MPa, preferably 2.0-6.0 MPa higher than the pressure in the first reactor, and the temperature in the second reactor is at least 15 ℃ lower, preferably 5-40 ℃ lower than the temperature in the first reactor.

10. The hydrogenation process according to claim 1, wherein the catalyst packed in the second reactor is a catalyst having a hydrodearomatic hydrocarbon function.

11. The hydrogenation process according to claim 1, wherein the S content of the refined naphtha is 600 μg/g or less, the olefin content is 1.0% or less, the aromatic hydrocarbon content in the refined diesel product is 5wt% or less, and the S content is 10 μg/g or less.

12. A hydrogenation reaction system for a hydrogenation process according to claim 1, comprising a coking unit and a hydrogenation unit, wherein the coking unit comprises a coking tower and a fractionating tower, the hydrogenation unit comprises a first reactor, namely 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 fractionating tower is connected with a raw material inlet of the first reactor.