CN111321006A

CN111321006A - Step-by-step diesel oil hydrogenation method based on different types of aromatic hydrocarbon compositions

Info

Publication number: CN111321006A
Application number: CN201811522329.9A
Authority: CN
Inventors: 刘涛; 李宝忠; 徐大海; 宣根海; 王晶晶; 陈�光; 郭蓉
Original assignee: China Petroleum and Chemical Corp; Sinopec Dalian Research Institute of Petroleum and Petrochemicals
Current assignee: Sinopec Dalian Petrochemical Research Institute Co ltd; China Petroleum and Chemical Corp
Priority date: 2018-12-13
Filing date: 2018-12-13
Publication date: 2020-06-23
Anticipated expiration: 2038-12-13
Also published as: CN111321006B

Abstract

The invention discloses a step-by-step diesel oil hydrogenation method based on different types of aromatic hydrocarbon, which comprises the following steps: (a) fractionating the diesel raw oil to obtain three fractions, namely light fraction, medium fraction and heavy fraction; (b) mixing the heavy fraction diesel raw oil obtained in the step (a) with recycle hydrogen, and then entering a first hydrogenation reaction zone; (c) mixing the resultant flow obtained in the step (b), the middle distillate raw oil obtained in the step (a) and hydrogen, and then entering a second hydrogenation reaction zone; (d) mixing the material flow obtained in the step (c), the light fraction diesel oil raw oil obtained in the step (a) and hydrogen, and then entering a third hydrogenation reaction zone; (e) separating the resultant stream obtained in step (d) into a gas phase and a liquid phase, and subjecting the liquid phase to a low pressure separator and fractionation or stripping to obtain a gas, naphtha and diesel product. The method obviously reduces the average reaction temperature, and generally shows that the investment cost of the overall equipment is reduced and the operation period of the device is prolonged.

Description

Step-by-step diesel oil hydrogenation method based on different types of aromatic hydrocarbon compositions

Technical Field

The invention belongs to a hydrogenation process of an oil refining technology, relates to a step-by-step diesel oil hydrogenation method based on different types of aromatic hydrocarbon compositions, and particularly relates to a low-energy-consumption hydrogenation process technology for producing high-quality diesel oil products by performing step-by-step hydrogenation on diesel oil fractions based on different types of aromatic hydrocarbon compositions.

Background

Energy currently worldwide is derived primarily from fossil energy sources, with petroleum being the most prominent source of motor fuel. As the world economy continues to evolve, environmental regulations become more stringent requiring the production of large quantities of light, clean motor fuels, which require improvements and modifications to existing refinery technologies. The quality requirements of diesel oil as important motor fuel are more and more high, and the indexes such as sulfur content, density, cetane number, polycyclic aromatic hydrocarbon and the like are strictly limited.

The diesel hydrogenation technology is the most important means for improving the quality of diesel products. CN102465021A discloses a diesel oil combined hydrogenation process method, in which a diesel oil fraction is cut into light and heavy components, wherein the light component is processed by a liquid phase hydrogenation method, and the heavy component is processed by a gas phase circulating hydrogenation reaction. CN102851067A discloses a hydrogenation method for producing ultra-low sulfur clean diesel, diesel fraction cutting 280～340 light and heavy components, wherein the heavy component is subjected to countercurrent hydrogenation, and the light component is subjected to conventional gas phase circulating hydrogenation. CN102051219A discloses a hydrotreating method of diesel fraction, wherein the diesel fraction is fractionated to obtain light and heavy fractions, the light fraction is hydrogenated conventionally, and the heavy fraction is hydrogenated by using a bulk phase catalyst. CN200410050729.6 discloses a diesel fraction hydrotreating method, which comprises fractionating raw oil into light fraction and heavy fraction, feeding the light fraction into a first reaction zone to contact with a hydrofining catalyst, feeding the heavy fraction into a second reaction zone to contact with the hydrofining catalyst, and mixing the liquid phase products of the two reaction zones to obtain refined distillate, wherein the catalysts and/or process conditions used in the two reaction zones are different. In the above method, the diesel oil is divided into light diesel oil and heavy diesel oilThe two fractions are respectively subjected to different hydrogenation processes, and the two hydrogenation products are mixed to obtain the diesel oil product without a common hydrogenation part. CN 102041071A discloses a diesel oil hydrogenation process, fractionating a diesel oil raw material into a light component and a heavy component, subjecting the heavy component to hydrofining reaction in a first reaction zone under hydrofining conditions, mixing the effluent of the first reaction zone and the light component, introducing the mixture into a second reaction zone for hydrofining reaction, and separating the effluent of the second reaction zone. In the method, heavy components are subjected to hydrodenitrogenation and hydrodearomatization reaction through a hydrogenation catalyst bed layer containing Mo-Ni or Mo-Ni-Co operated at low temperature, only a small part of hydrodesulfurization reaction is carried out, and hydrogenation reaction material flow and light components are mixed and subjected to hydrodesulfurization through a hydrogenation catalyst bed layer containing Mo-Co operated at high temperature. The method can cause insufficient aromatic hydrocarbon saturation due to thermodynamic limitation caused by overhigh reaction temperature when the device is processed to the final stage, the quality of a diesel oil product is influenced, and the sulfur content can not meet the index requirement.

In summary, in the prior art, as the diesel raw material contains S which exists in the form of sulfide and generates hydrogen sulfide through hydrogenation reaction, and the sulfide reacts easily after contacting with the catalyst bed, the concentration of hydrogen sulfide rises immediately, the hydrogen sulfide content is higher and higher along with the material flowing direction, the hydrogen sulfide has obvious inhibiting effect on hydrodesulfurization, and especially has more obvious inhibiting effect on deep hydrodesulfurization of complex sulfide which is difficult to react, but the prior art has a fresh report on the technology for reducing the influence of hydrogen sulfide generated in the reaction process, in addition, the diesel hydrogenation operation is usually carried out under medium pressure conditions, when the final-stage operation is carried out, the aromatic hydrocarbon hydrogenation saturation is limited by thermodynamics along with the rise of the reaction temperature, so that the aromatic hydrocarbon content in the product is higher, and even the sulfur content does not meet the index requirement of ultra-low sulfur diesel.

Disclosure of Invention

Aiming at the defects of the prior art, the invention provides a step-by-step diesel oil hydrogenation process based on different types of aromatic hydrocarbon compositions. The method can fractionate the raw oil into light, medium and heavy fractions, control the reaction temperature and reduce the influence of hydrogen sulfide concentration on hydrodesulfurization of components difficult to react, obviously reduce the average reaction temperature when the total airspeed is the same to achieve the same target product quality, and generally show that the total equipment investment cost is reduced or the operation period of the device is prolonged.

The invention relates to a step-by-step diesel oil hydrogenation method based on different types of aromatic hydrocarbons, which comprises the following steps:

(a) fractionating the diesel raw oil to obtain three fractions, namely light fraction, medium fraction and heavy fraction;

(b) mixing the heavy fraction diesel raw oil obtained in the step (a) with recycle hydrogen, then feeding the mixture into a hydrofining catalyst bed in a first hydrogenation reaction zone, reacting under the hydrogenation operation condition to generate a material flow generated in the first hydrogenation reaction zone, wherein the sulfur content in the liquid material flow is not more than 35 mug/g;

(c) mixing the generated material flow of the first hydrogenation reaction zone obtained in the step (b), the middle distillate raw oil obtained in the step (a) and hydrogen, then feeding the mixture into a hydrofining catalyst bed layer in a second hydrogenation reaction zone, and reacting under the hydrogenation operation condition to generate a material flow of the second hydrogenation reaction zone, wherein the sulfur content in the liquid material flow is not more than 35 [ mu ] g/g;

(d) mixing the material flow generated in the second hydrogenation reaction zone obtained in the step (c) with the light fraction diesel oil raw material oil and hydrogen obtained in the step (a), then feeding the mixture into a hydrofining catalyst bed layer in a third hydrogenation reaction zone, and reacting under the hydrogenation operation condition to generate a material flow generated in the third hydrogenation reaction zone;

(e) and (d) separating the generated material flow of the third hydrogenation reaction zone obtained in the step (d) into a gas phase and a liquid phase in a high-pressure separator, recycling the hydrogen-rich gas obtained by separation directly or after removing hydrogen sulfide, and obtaining gas, naphtha and diesel oil products after the liquid phase obtained by separation is subjected to low-pressure separator and fractionation or stripping.

In the method, the sulfur content in the liquid stream of the first hydrogenation reaction zone and/or the second hydrogenation reaction zone is not more than 35 mug/g, preferably 10 mug/g-30 mug/g.

In the method, the ratio of the increase of the hydrogen sulfide of the first reaction zone (the content difference between the outlet of the first reaction zone and the inlet of the first reaction zone) to the increase of the total hydrogen sulfide (the content difference between the outlet of the third reaction zone and the inlet of the first reaction zone) is 23-50%, preferably 26-48%.

In the above method, it is further preferable to adopt one of the following schemes:

1) heavy fraction firstly enters a first hydrogenation reaction zone and passes through an upper part of a hydrofining catalyst bed layer, one part of generated material flow enters a gas-liquid separator to be separated into a gas phase and a liquid phase, the other part of generated material flow and the liquid phase continuously pass through a lower part of the hydrofining catalyst bed layer, the generated material flow and middle fraction are mixed and enter a second hydrogenation reaction zone and are hydrofined through the hydrofining catalyst bed layer, and the generated material flow, light fraction and gas phase are mixed and then continuously enter a third reaction zone and are hydrofined through the hydrofining catalyst bed layer;

2) heavy fraction firstly enters a first hydrogenation reaction zone and passes through an upper part hydrofining catalyst bed layer, one part of product flow enters a gas-liquid separator to be separated into a gas phase and a liquid phase, the other part of product flow and the liquid phase continuously pass through a lower part of hydrofining catalyst bed layer, the product flow and middle fraction are mixed and enter an upper part of hydrofining catalyst bed layer of a second hydrogenation reaction zone, one part of product flow enters the gas-liquid separator to be separated into the gas phase and the liquid phase, the other part of product flow and the liquid phase continuously pass through the lower part of hydrofining catalyst bed layer to be hydrofined, and the product flow, light fraction, the gas phase flow of the first reaction zone and the gas phase flow of the second reaction zone are mixed and then continuously enter a third reaction zone to be hydrofined through the hydrof;

3) heavy fraction firstly enters a first reaction zone to be subjected to hydrofining and hydro-upgrading through a hydrofining catalyst bed layer and a hydro-upgrading catalyst bed layer, a generated material flow and middle fraction are mixed and enter a second reaction zone to be subjected to hydrofining and hydro-upgrading through the hydrofining catalyst bed layer and the hydro-upgrading catalyst bed layer, and the generated material flow and light fraction are mixed and then continuously enter a third reaction zone to be subjected to hydrofining through the hydro-upgrading catalyst bed layer;

4) heavy fraction firstly enters a first hydrogenation reaction zone and passes through an upper part of a hydrofining catalyst bed layer, one part of product flow enters a gas-liquid separator to be separated into a gas phase and a liquid phase, the other part of product flow and the liquid phase continue to pass through a lower part of the hydrofining catalyst bed layer to be subjected to hydrofining and hydro-upgrading, the product flow and middle fraction are mixed and enter a second reaction zone and pass through the hydrofining catalyst bed layer and the hydro-upgrading catalyst bed layer to be subjected to hydrofining and hydro-upgrading, and the product flow, light fraction and the gas phase flow of the first reaction zone are mixed and then continue to enter a third reaction zone and pass through the hydrofining catalyst bed layer to be;

5) heavy fraction enters a first hydrogenation reaction zone and passes through an upper part hydrofining catalyst bed layer, one part of product flow enters a gas-liquid separator to be separated into a gas phase and a liquid phase, the other part of product flow and the liquid phase continue to pass through a lower part of hydrofining catalyst bed layer to be hydrofined and hydro-modified, the product flow and middle fraction are mixed and enter an upper part of hydrofining catalyst bed layer of a second reaction zone, one part of product flow enters a gas-liquid separator to be separated into a gas phase and a liquid phase, the other part of product flow and the liquid phase continue to pass through a lower part of hydrofining catalyst bed layer and a hydro-modifying catalyst bed layer to be hydrofined and hydro-modified, the product flow, light fraction and the gas phase material flow of, after mixing the gas phase material flows in the second reaction zone, continuously entering a third reaction zone and carrying out hydrofining through a hydrofining catalyst bed layer;

6) heavy fraction firstly enters a first reaction zone, is subjected to hydrofining and hydro-upgrading through a hydrofining catalyst bed and a hydro-upgrading catalyst bed, a generated material flow and middle fraction are mixed and enter a second reaction zone, are subjected to hydrofining through the hydrofining catalyst bed, and the generated material flow and light fraction are mixed and then continuously enter a third reaction zone, and are subjected to hydrofining through the hydrofining catalyst bed;

7) heavy fraction firstly enters a first hydrogenation reaction zone and passes through an upper part of a hydrofining catalyst bed layer, one part of product flow enters a gas-liquid separator to be separated into a gas phase and a liquid phase, the other part of product flow and the liquid phase continue to pass through a lower part of the hydrofining catalyst bed layer to be subjected to hydrofining and hydro-upgrading, the product flow and middle fraction are mixed and enter a second reaction zone and are subjected to hydrofining through the hydrofining catalyst bed layer, the product flow is mixed with light fraction and the gas phase flow of the first reaction zone and then continues to enter a third reaction zone and are subjected to hydro-refining through the hydrofining catalyst bed layer;

8) heavy fraction firstly enters a first hydrogenation reaction zone and passes through an upper part hydrofining catalyst bed layer, one part of product flow enters a gas-liquid separator to be separated into a gas phase and a liquid phase, the other part of product flow and the liquid phase continue to pass through a lower part of hydrofining catalyst bed layer to be subjected to hydrofining and hydro-upgrading, the product flow and middle fraction enter an upper part of hydrofining catalyst bed layer of a second reaction zone, one part of product flow enters a gas-liquid separator to be separated into a gas phase and a liquid phase, the other part of product flow and the liquid phase continue to pass through a lower part of hydrofining catalyst bed layer to be subjected to hydrofining, and the product flow, light fraction, the gas phase material flow of the first reaction zone and the gas phase material flow of the second reaction zone are mixed and then continue to enter a;

9) heavy fraction firstly enters a first reaction zone and is hydrofined through a hydrofining catalyst bed, a generated material flow and middle fraction are mixed and enter a second reaction zone and are hydrofined and hydroupgraded through a hydrofining catalyst bed and a hydro-upgrading catalyst bed, and the generated material flow and light fraction are mixed and then continuously enter a third reaction zone and are hydrofined through a hydrofining catalyst bed;

10) heavy fraction firstly enters a first hydrogenation reaction zone and passes through an upper part of a hydrofining catalyst bed layer, one part of product flow enters a gas-liquid separator to be separated into a gas phase and a liquid phase, the other part of product flow and the liquid phase continue to pass through a lower part of the hydrofining catalyst bed layer for hydrofining, the product flow and middle fraction are mixed and enter a second reaction zone and pass through the hydrofining catalyst bed layer and the hydrofining catalyst bed layer for hydrofining and hydro-upgrading, and the product flow, light fraction and the gas phase flow in the first reaction zone are mixed and then continue to enter a third reaction zone and pass through the hydrofining catalyst bed layer for hydrofining;

11) heavy fraction firstly enters a first hydrogenation reaction zone and passes through an upper part hydrofining catalyst bed layer, one part of product flow enters a gas-liquid separator to be separated into a gas phase and a liquid phase, the other part of product flow and the liquid phase continue to pass through a lower part of hydrofining catalyst bed layer to be hydrofined, the product flow and middle fraction are mixed and enter an upper part of hydrofining catalyst bed layer of a second reaction zone, one part of product flow enters a gas-liquid separator to be separated into a gas phase and a liquid phase, the other part of product flow and the liquid phase continue to pass through a lower part of hydrofining catalyst bed layer and a hydrofining catalyst bed layer to be hydrofined and hydro-modified, and the product flow, light fraction, the gas phase flow of the first reaction zone and the gas phase flow of the second reaction zone are mixed and then continue to enter a third reaction zone.

In the method, the used diesel raw oil can comprise one or more of straight-run diesel, catalytic diesel, coking diesel, thermal cracking diesel, visbreaking diesel, synthetic diesel, coal tar diesel fraction, coal direct liquefaction diesel, shale oil diesel and other diesel fractions.

In the method, the fractionation point temperature of the light fraction and the middle fraction obtained from the raw diesel oil is 200-270 ℃, and preferably 210-265 ℃. Based on the selected cut point temperature, the weight content of monocyclic aromatics in the light fraction is greater than 92%, preferably greater than 95%, of the weight content of total monocyclic aromatics in the diesel feedstock.

In the method, the fractionation point temperature of the middle distillate and the heavy distillate obtained from the raw diesel oil is 260-320 ℃, and preferably 270-315 ℃. Based on the selected cut point temperature, the weight content of tricyclic and higher tricyclic aromatics in the heavy fraction is greater than 92%, preferably greater than 95%, of the weight content of total tricyclic and higher tricyclic aromatics in the diesel feedstock.

In the method, the weight ratio of the light fraction to the raw oil obtained in the fractionation temperature range is 20-40%, the weight ratio of the medium fraction to the raw oil is 25-45%, and the weight ratio of the heavy fraction to the raw oil is 25-45%.

In the above method, the first hydrogenation reaction zone may be one reactor or may be composed of a plurality of reactors, and if a plurality of reactors are selected, it is preferable that 2 to 4 reactors are selected. The reactor can be internally provided withOne catalyst bed layer is arranged, a plurality of catalyst bed layers can be arranged, and if a plurality of catalyst bed layers are selected, the number of catalyst bed layers is preferably 2-5. The amount of cold hydrogen between the catalyst beds, or between the reactors, is determined by the reactant stream flow rate and the temperature difference that needs to be reduced. The operating conditions of the first hydrogenation reaction zone are that the reaction pressure is 3.0-14.0 MPa, and the volume space velocity of the heavy fraction diesel oil raw material oil is 0.2h^-1~10.0h^-1The average reaction temperature is 220-450 ℃, and the volume ratio of hydrogen to oil is 200: 1-2000: 1; the preferable operation conditions are that the reaction pressure is 4.0 MPa-13.0 MPa, the volume space velocity of the heavy fraction raw oil is 0.3h < -1 > -8.0 h < -1 >, the average reaction temperature is 230 ℃ to 440 ℃, and the volume ratio of hydrogen to oil is 300: 1-1800: 1.

In the above method, for a preferred embodiment, the hydro-upgrading catalyst used in the first hydrogenation reaction zone is a conventional diesel hydro-upgrading catalyst, and generally uses a group vib and/or group viii metal as an active component, the group vib metal is generally Mo and/or W, and the group viii metal is generally Co and/or Ni. The catalyst carrier contains one or more of alumina, siliceous alumina and molecular sieve, preferably containing molecular sieve; the molecular sieve can be a Y-type molecular sieve. Based on the weight of the catalyst, the content of the VIB group metal is 10-35 wt% calculated by oxide, the content of the VIII group metal is 3-15 wt% calculated by oxide, the content of the molecular sieve is 5-40 wt%, and the content of the alumina is 10-80 wt%; its specific surface area is 100m²/g～650m²The pore volume is 0.15mL/g to 0.50 mL/g. The main catalysts comprise 3963, FC-18, FC-32 catalysts and the like which are developed by the petrochemical research institute. For the hydrogenation modification catalyst, certain hydrogenation activity and certain cracking activity are required, and both hydrogenation saturation of olefin and aromatic hydrocarbon in diesel oil fraction and ring-opening reaction of saturated aromatic hydrocarbon are required. The operation conditions for the hydro-upgrading of the first hydrogenation reaction zone can adopt conventional operation conditions, and generally comprise: the reaction pressure is 3.0MPa to 13.0MPa, the reaction temperature is 300 ℃ to 433 ℃, and the liquid hourly volume space velocity is 0.3h^-1～15.0h^-1The volume ratio of hydrogen to oil is 100: 1-2000: 1

In the above process, the second hydrogenation reaction zone may be a reactor or a reactorThe reactor is composed of a plurality of reactors, and if a plurality of reactors are selected, the number of the reactors is preferably 2-4. One catalyst bed layer or a plurality of catalyst bed layers can be arranged in the reactor, and preferably 2-5 catalyst bed layers are selected if a plurality of catalyst bed layers are selected. The amount of cold hydrogen between the catalyst beds, or between the reactors, is determined by the flow rate of the stream and the temperature difference that needs to be reduced. The operating conditions of the second hydrogenation reaction zone are that the reaction pressure is 3.0 MPa-14.0 MPa, and the total volume space velocity of the light fraction and heavy fraction diesel raw oil is 0.2h^-1~12.0h^-1The average reaction temperature is 210-448 ℃, and the volume ratio of hydrogen to oil (the volume ratio of hydrogen to all fresh raw oil) is 150: 1-1700: 1; the preferable operation conditions are that the reaction pressure is 4.0 MPa-13.0 MPa, and the total volume airspeed of the diesel raw oil is 0.3h^-1~10.0h^-1The average reaction temperature is 220-438 ℃, and the volume ratio of hydrogen to oil (the volume ratio of hydrogen to all fresh raw oil) is 200: 1-1300: 1.

In the above method, for a preferred embodiment, the hydro-upgrading catalyst used in the second hydrogenation reaction zone is a conventional diesel hydro-upgrading catalyst, and generally uses a group vib and/or group viii metal as an active component, the group vib metal is generally Mo and/or W, and the group viii metal is generally Co and/or Ni. The catalyst carrier contains one or more of alumina, siliceous alumina and molecular sieve, preferably containing molecular sieve; the molecular sieve can be a Y-type molecular sieve. Based on the weight of the catalyst, the content of the VIB group metal is 10-35 wt% calculated by oxide, the content of the VIII group metal is 3-15 wt% calculated by oxide, the content of the molecular sieve is 5-40 wt%, and the content of the alumina is 10-80 wt%; its specific surface area is 100m²/g～650m²The pore volume is 0.15mL/g to 0.50 mL/g. The main catalysts comprise 3963, FC-18, FC-32 catalysts and the like which are developed by the petrochemical research institute. For the hydrogenation modification catalyst, certain hydrogenation activity and certain cracking activity are required, and both hydrogenation saturation of olefin and aromatic hydrocarbon in diesel oil fraction and ring-opening reaction of saturated aromatic hydrocarbon are required. The operation conditions of the second hydrogenation reaction zone for hydro-upgrading can adopt conventional operation conditions, and generally comprise: the reaction pressure is 3.0MPa to 13.0MPa, and the reaction temperature is 300 DEG CLiquid hourly volume space velocity of 0.3h at 433 DEG C^-1～15.0h^-1The volume ratio of hydrogen to oil is 100: 1-2000: 1

In the method, the sulfur content in the liquid stream in the second hydrogenation reaction zone is not more than 35 mug/g, preferably 10 mug/g-30 mug/g.

In the method, the average reaction temperature of the first reaction zone can be higher than the average reaction temperature of the second reaction zone, can be the same or lower, and is preferably higher, and the difference between the two average reaction temperatures is 2-35 ℃, and is preferably 3-30 ℃.

In the above method, the third hydrogenation reaction zone may be one reactor or may be composed of a plurality of reactors, and if a plurality of reactors are selected, it is preferable that 2 to 4 reactors are selected. One catalyst bed layer or a plurality of catalyst bed layers can be arranged in the reactor, and preferably 2-5 catalyst bed layers are selected if a plurality of catalyst bed layers are selected. The amount of cold hydrogen between the catalyst beds, or between the reactors, is determined by the flow rate of the stream and the temperature difference that needs to be reduced. The operating conditions of the second hydrogenation reaction zone are that the reaction pressure is 3.0 MPa-14.0 MPa, and the total volume space velocity of the light fraction and heavy fraction diesel raw oil is 0.2h^-1~15.0h^-1The average reaction temperature is 200-445 ℃, and the volume ratio of hydrogen to oil (the volume ratio of hydrogen to all fresh raw oil) is 150: 1-1500: 1; the preferable operation conditions are that the reaction pressure is 4.0 MPa-13.0 MPa, and the total volume airspeed of the diesel raw oil is 0.3h^-1~12.0h^-1The average reaction temperature is 210-435 ℃, and the volume ratio of hydrogen to oil (the volume ratio of hydrogen to all fresh raw oil) is 200: 1-1200: 1.

In the method, the average reaction temperature of the second reaction zone can be higher than the average reaction temperature of the third reaction zone, or can be the same or lower, preferably, the average reaction temperature of the second reaction zone is higher, and the difference between the two average reaction temperatures is 3-40 ℃, preferably 5-35 ℃.

In the method, the hydrogenation reaction is an exothermic reaction, and the reaction temperature in the reaction zone tends to increase, namely the outlet temperature of the first reaction zone is higher than the inlet temperature of the first reaction zone, the outlet temperature of the second reaction zone is higher than the inlet reaction temperature, and the outlet temperature of the third reaction zone is higher than the inlet reaction temperature. Controlling the outlet temperature of the first reaction zone to be higher than the inlet temperature of the second reaction zone, wherein the difference between the two temperatures is 5-80 ℃, preferably 15-70 ℃, and the larger the difference between the two temperatures is, the smaller the temperature required by the middle distillate is, and the lower the overall energy consumption of the device is. The difference between the two temperatures is adjusted by the weight and temperature of the middle distillate fraction, and recycle hydrogen can be injected to assist adjustment. Wherein the temperature of the middle distillate can be obtained by heat exchange or heating. The outlet temperature of the second reaction zone is controlled to be higher than the inlet temperature of the third reaction zone, and the difference between the two temperatures is 5-80 ℃, and preferably 15-70 ℃. The greater the difference between these two temperatures, the lower the temperature to which the light fraction needs to be increased and the lower the overall energy consumption of the plant. The difference between these two temperatures is regulated by the weight and temperature of the light ends and may be assisted by the injection of recycle hydrogen. Wherein the temperature of the light fraction can be obtained by means of heat exchange or heating.

In the above process, in order to reduce the effect of hydrogen sulfide on hydrodesulfurization, particularly heavy fraction hydrodesulfurization, it is necessary to control the total amount of sulfur entering the first hydrogenation reaction zone. The ratio of the hydrogen sulfide increase value of the first reaction zone (the content difference between the outlet of the first reaction zone and the inlet of the first reaction zone) to the total hydrogen sulfide increase value (the content difference between the outlet of the third reaction zone and the inlet of the first reaction zone) is controlled to be 23-50%, preferably 26-48%. In addition to adequately reducing the effect of hydrogen sulfide on the hydrodesulfurization of the middle distillate, it is desirable to control the total amount of sulfur entering the second hydrogenation reaction zone. The ratio of the hydrogen sulfide increase value of the second reaction zone (the content difference between the outlet of the second reaction zone and the inlet of the second reaction zone) to the total hydrogen sulfide increase value (the content difference between the outlet of the third reaction zone and the inlet of the first reaction zone) is controlled to be 23-48%, preferably 26-46%.

In the method, the hydrogenation active component in the hydrofining catalyst is one or more of Co, Mo, W and Ni, the weight content of the hydrogenation active component is 5-70% in terms of oxide, the carrier of the hydrofining catalyst is generally alumina, amorphous silicon aluminum, silicon oxide, titanium oxide and the like, and other auxiliary agents such as P, Si, B, Ti, Zr and the like can be simultaneously contained. The catalyst may be used commercially or may be prepared by methods known in the art. The hydrogenation active component is a catalyst in an oxidation state, and is subjected to conventional vulcanization treatment before use, so that the hydrogenation active component is converted into a vulcanization state. The commercial hydrogenation catalysts mainly comprise hydrogenation catalysts such as 3936, CH-20, FF-14, FF-26, FH-5A, FH-98, FH-DS series, FH-UDS series and the like developed by the Fushu petrochemical research institute (FRIPP), hydrogenation catalysts such as HR-416, HR-448 and the like of IFP company, hydrogenation catalysts such as HC-P, HC-K UF-210/220 of UOP company, and hydrogenation catalysts such as KF-752, KF-840, KF-848, KF-901, KF-907 and the like of AKZO company. The hydrogenation catalyst used in the first hydrogenation reaction zone, the hydrogenation catalyst used in the second hydrogenation reaction zone and the hydrogenation catalyst used in the third hydrogenation reaction zone can be the same or different.

In the above method, in a preferred embodiment, the gas-liquid separator is a device disposed between the first hydrofining reaction zone and the hydrofining catalyst bed. The gas-liquid separator at least comprises a reaction material flow inlet, a liquid phase conduit, a gas phase conduit and the like, wherein the gas phase conduit extracts a gas phase obtained by separation out of the first hydrofining reaction zone, and the liquid phase conduit continuously introduces a liquid phase obtained by separation downwards into the hydrofining catalyst bed layer at the lower part.

In the above method, in a preferred embodiment, the gas-phase stream extracted by the gas-liquid separator accounts for 10v% to 50% of the volume of the gas-phase stream at the inlet of each reaction zone.

In the method, the hydrogenation reaction effluent is separated by a high-pressure separator and/or a low-pressure separator. The high-pressure separator is a conventional gas-liquid separator. And the hydrogenation reaction material flow is separated in a high-pressure separator to obtain high-pressure hydrogen-rich gas and liquid. The low-pressure separator is a conventional gas-liquid separator. The liquid obtained by separation in the high-pressure separator is separated in the high-low pressure separator to obtain gas and liquid.

In the method, the high-pressure hydrogen-rich gas obtained by the separation of the high-pressure separator can be directly recycled by a circulating compressor or recycled after hydrogen sulfide is removed by a desulfurizing tower.

In the above method, the fresh hydrogen of the apparatus is pressurized by the fresh hydrogen compressor and then mixed with the recycle hydrogen at the inlet or outlet of the recycle hydrogen compressor, preferably at the outlet.

In the method, the fractionating system used for fractionating comprises a stripping tower and/or a fractionating tower. And the liquid obtained by separation in the low-pressure separator is subjected to steam stripping and/or fractionation in a fractionation system to obtain a naphtha product and a diesel product.

In the above method, the gas separator used for gas separation is a conventional separator. And after the gas obtained by separation in the low-pressure separator is mixed, the gas is separated in a gas separator to obtain hydrogen, dry gas, liquefied gas and the like, and if a liquid product exists, the gas directly enters a stripping tower and/or a fractionating tower.

Research finds that sulfur in raw oil in a diesel hydrogenation process generates hydrogen sulfide through hydrogenation reaction, the hydrogen sulfide is an important factor for restricting hydrodesulfurization, particularly deep desulfurization, when full fractions of a diesel raw material pass through a catalyst bed layer together, easily-reacted sulfides preferentially react, difficultly-reacted complex sulfides react finally, the concentration of the hydrogen sulfide gradually rises along with the flowing direction of a reactant flow, the difficulty of the deep hydrodesulfurization reaction is further increased under the double action, and the specification requirement of a target product is met by increasing the reaction temperature and reducing the volume airspeed, so that the operation period is shortened, or the device is operated in a reduced amount mode. The research result shows that the diesel fraction is fractionated to obtain three fractions of light fraction, medium fraction and heavy fraction, the heavy fraction with complex sulfide content difficult to react firstly enters the catalyst bed layer of the first reaction zone, the reacted material and the medium fraction with high sulfide content moderate to react are mixed and enter the catalyst bed layer of the second reaction zone, then the reacted material is continuously mixed with the light fraction with high sulfide content easy to react and enters the catalyst bed layer of the third reaction zone, the hydrogen sulfide concentration and the difficulty degree of the reaction sulfide are optimally combined, and the hydrodesulfurization efficiency is improved on the premise of not influencing the quality of the diesel product, the running period is generally prolonged, or the equipment investment is reduced by controlling the difference of the average reaction temperature of the three reaction zones and the difference of the hydrogen-oil volume ratio of the three reaction zones.

In the prior art, the diesel raw material can be hydrogenated by a fixed bed to produce clean diesel, in the whole hydrogenation reaction process, particularly, the diesel raw oil is in the initial period of contact with a catalyst, wherein the easily-reacted sulfide in the light fraction is firstly subjected to hydrogenation reaction, so that the concentration of hydrogen sulfide is rapidly increased, namely, the concentration of hydrogen sulfide in the whole reaction process is very high and gradually increases along with the flowing direction of material flow, when the subsequent difficultly-reacted sulfide is subjected to hydrogenation reaction, particularly, the most difficult reactant mainly comprising 4, 6-dimethyldibenzothiophene is frequently generated at the extreme end of all catalysts during deep hydrodesulfurization, at this time, the concentration of hydrogen sulfide almost reaches the maximum, the difficulty of deep hydrodesulfurization is obviously increased, and therefore, the aim requirements are met at higher reaction temperature, lower volume space velocity or higher reaction pressure, higher reaction temperatures result in higher aromatics content in the product, especially polycyclic aromatics content, due to thermodynamic equilibrium limitations. The invention fully utilizes the complex sulfide which is most difficult to react at low hydrogen sulfide concentration to react first, the complex sulfide which is relatively difficult to react at medium hydrogen sulfide concentration to react, and the sulfide which is relatively easy to react at high hydrogen sulfide concentration to react last, thereby reducing the influence of hydrogen sulfide on the hydrodesulfurization reaction on the whole, and controlling the average reaction temperature of the first reaction zone to be higher than the average reaction temperature of the second reaction zone, and controlling the average reaction temperature of the second reaction zone to be higher than the average reaction temperature of the third reaction zone so as to realize better hydrodesulfurization and better aromatic hydrocarbon hydrogenation saturation. The diesel raw oil is firstly fractionated into light fraction, medium fraction and heavy fraction, the heavy fraction is firstly contacted with a catalyst, the most difficult-to-react sulfide mainly containing three or more aromatic rings is subjected to hydrodesulfurization reaction in a lower hydrogen sulfide atmosphere of a first reaction zone, the increase of the hydrogen sulfide concentration in the most initial catalyst bed layers is slower because of higher hydrodesulfurization difficulty, and the reaction performance of the catalyst is correspondingly enhanced because of no competitive adsorption of the sulfide of the medium fraction and the light fraction even if the volume space velocity of the heavy fraction is the same. The hydrogenation reaction material flow of the heavy fraction is continuously mixed with the middle fraction and then contacted with the catalyst of the second reaction area, the sulfide which is hard to react and mainly takes two aromatic ring structures is continuously subjected to hydrodesulfurization reaction under the hydrogenation condition, then the reaction material flow of the second reaction area is continuously mixed with the light fraction and then contacted with the catalyst of the third reaction area, the sulfide which is easy to react and mainly takes monocyclic aromatic hydrocarbon and paraffin hydrocarbon is continuously subjected to hydrodesulfurization reaction under the hydrogenation condition with higher hydrogen sulfide concentration, and the requirement of producing clean diesel oil products is completely met. The temperature of the second reaction zone is further lowered by 2-35 ℃ than that of the first reaction zone, the temperature of the third reaction zone is lowered by 3-40 ℃ than that of the second reaction zone, namely, the deep removal of heavy fraction sulfides is realized under the condition of higher reaction temperature in the first reaction zone, the hydrodearomatization of heavy fraction and the deep hydrodesulfurization and hydrodearomatization of middle fraction are realized under the condition of relatively lower reaction temperature in the second reaction zone, the hydrodearomatization of heavy fraction and middle fraction and the deep hydrodesulfurization and hydrodearomatization of light fraction are realized under the condition of lower reaction temperature in the third reaction zone, namely, the purposes of deep desulfurization and deep dearomatization are simultaneously achieved through different reaction conditions of the three reaction zones; the method for controlling the temperature difference between the two reaction zones is more, preferably, the heavy fraction raw oil meets the higher temperature requirement by using a heat exchange and heating method, the medium fraction raw oil meets the relatively higher temperature requirement by using a heat exchange and/or heating method, and the light fraction raw oil can meet the lower temperature requirement by using a heat exchange method only. In addition, because the raw oil in the first reaction zone is only heavy fraction raw oil, hydrogen required by hydrogenation reaction is relatively less, less recycle hydrogen can be used when the same hydrogen-oil volume ratio is reached, the operation cost is saved, or the larger hydrogen-oil volume ratio is provided under the condition of using the same recycle hydrogen, the advantages of relatively high hydrogen partial pressure and relatively lower hydrogen sulfide partial pressure in the surface environment of the catalyst pore channel are brought, and the hydrogenation reaction of sulfide which is most difficult to react and the hydrogenation saturation of polycyclic aromatic hydrocarbon are more facilitated. In the same way, because the raw oil in the second reaction zone is heavy fraction and medium fraction, the hydrogen-oil ratio lower than that of the first reaction zone can be used to meet the reaction requirement, and the lowest hydrogen-oil ratio can be used in the third reaction zone to meet the requirement of the final product, namely, the distribution of hydrogen and the reaction conditions of different reaction zones are optimized on the premise of meeting the requirement of the lowest hydrogen-oil ratio of the third reaction zone in the three reaction zones. The stepwise diesel hydrogenation process based on different types of aromatic hydrocarbons generally shows that the internal reaction environment is optimized and the average reaction temperature is reduced on the premise of not influencing the quality of diesel products, so that the operation period is prolonged, or the equipment investment is saved.

The invention further optimizes the use of the catalyst, namely a Co-containing hydrogenation catalyst mainly based on direct desulfurization activity is used in the first reaction zone and the second reaction zone, and a Mo-Ni hydrogenation catalyst mainly based on hydrogenation and dearomatization is used in the third hydrogenation reaction zone. The advantage of this option is that the removal of sulphur from the heavy fraction in the first reaction zone at low hydrogen sulphide concentration and the removal of sulphur from the medium fraction in the second reaction zone at medium hydrogen sulphide concentration allows the deep hydrodesulphurisation of the heavy and medium fractions to be achieved even at higher temperatures towards the end of the run, whereas in the third reaction zone, the temperatures are still relatively low towards the end of the run, completely satisfying the hydrodesulphurisation requirements and being less thermodynamically constrained, thus achieving longer line cycles or saving equipment investment.

In the preferred scheme of the method, part of the hydrogenation modified catalyst is continuously used after the heavy fraction is hydrofined, so that the difficulty of hydrodesulfurization can be reduced by utilizing the acidity in the modified catalyst, the outermost ring of the polycyclic aromatic hydrocarbon can be opened after being subjected to hydrogenation saturation, and the density of a diesel product and the cetane number can be further reduced on the basis of keeping the advantages of reducing energy consumption and overall equipment investment cost, prolonging the operation period of a device and the like. In the preferred scheme of the method, part of the hydrogenation modified catalyst is continuously used after the heavy middle distillate is hydrofined, so that the difficulty of hydrodesulfurization can be reduced by utilizing the acidity in the modified catalyst, and the outermost ring of the bicyclic aromatic hydrocarbon can be opened after being subjected to hydrogenation saturation, thereby further reducing the density of a diesel product and improving the cetane number.

In the preferred scheme of the method, part of the generated material flow after the heavy fraction passes through the part of the hydrofining on the first hydrogenation reaction zone is separated by the gas-liquid separator, and the separated gas phase rich in hydrogen sulfide is led out of the heavy fraction hydrofining reaction zone, so that the content of hydrogen sulfide in the system during deep desulfurization of the refractory sulfide is further reduced, and the deep hydrodesulfurization reaction of the refractory sulfide is facilitated. And (3) separating a part of the generated material flow after the middle fraction passes through the upper part of the second hydrogenation reaction zone and is subjected to hydrofining by a gas-liquid separator, and leading the separated gas phase rich in hydrogen sulfide out of the middle fraction hydrofining reaction zone, so that the hydrogen sulfide content in the system during deep desulfurization of the refractory sulfide is further reduced, and the deep hydrodesulfurization reaction of the refractory sulfide is facilitated. The gas phase rich in hydrogen sulfide does not have great influence on the hydrodesulfurization of the light fraction introduced into the third hydrogenation reaction zone, and can bring certain heat, and the first reaction zone, the second reaction zone and the third reaction zone form coupling operation, so that the energy consumption of the device is further reduced.

The invention further optimizes and uses the catalysts, namely a Mo-Ni type hydrogenation catalyst mainly used for hydrogenation and dearomatization and a Mo-Co type or Mo-Ni-Co type hydrogenation catalyst mainly used for direct desulfurization activity are used in a first reaction zone in a grading manner, the volume ratio of the two catalysts is 20: 80-45: 55, a Mo-Ni type hydrogenation catalyst mainly used for hydrogenation and dearomatization and a Mo-Co type or Mo-Ni-Co type hydrogenation catalyst mainly used for direct desulfurization activity are used in a grading manner in a second reaction zone, the volume ratio of the two catalysts is 20: 80-35: 65, and a Mo-Ni type hydrogenation catalyst mainly used for hydrogenation and dearomatization is used in a third hydrogenation reaction zone. The advantage of such selection is that the sulfur in the heavy fraction and the polycyclic aromatic hydrocarbon are removed under the condition of low hydrogen sulfide concentration in the first reaction zone and the sulfur in the middle fraction is removed under the condition of medium hydrogen sulfide concentration in the second reaction zone on the premise of keeping the advantages, the requirement of deep hydrodesulfurization of the heavy fraction and the middle fraction and the partial hydrogenation saturation of the polycyclic aromatic hydrocarbon can be achieved even if the temperature is higher during the last stage operation, and the temperature is still relatively lower during the last stage operation in the third reaction zone, so that the requirement of hydrodesulfurization can be completely met, the thermodynamic limitation is smaller, and the polycyclic aromatic hydrocarbon content can meet the requirement, thereby realizing the purpose of prolonging the line period or saving the equipment investment.

Drawings

FIG. 1 is a schematic flow diagram of a hydrogenation process according to the present invention.

FIG. 2 is another schematic flow diagram of the hydrogenation process of the present invention. FIG. 2 differs from FIG. 1 in the addition of a first reaction zone gas-liquid separator and a first reaction zone withdrawal of the gas phase.

Wherein in fig. 1 and 2: 1-heavy fraction raw oil, 2-first raw oil pump, 3-middle fraction raw oil, 4-second raw oil pump, 5-light fraction raw oil, 6-second raw oil pump, 7-first hydrogenation reactor, 8-second hydrogenation reactor, 9-third hydrogenation reactor, 10-first hydrogenation reaction effluent, 11-second hydrogenation reaction effluent, 12-third hydrogenation reaction effluent, 13-high pressure separator, 14-high pressure separator gas, 15-stripping/fractionating system, 16-naphtha, 17-diesel, 18-recycle hydrogen compressor, 19-new hydrogen, 20-raw oil supply unit, 21-first reaction zone gas-liquid separator, 22-first reaction zone gas phase extraction, 23-second reaction zone gas-liquid separator, 24-the second reaction zone withdraws the gaseous phase.

Detailed Description

The flow and effect of the step-by-step diesel hydrogenation process based on different types of aromatics according to the present invention will be further described with reference to the following examples, wherein the first hydrogenation zone is a hydrogenation reactor, the second hydrogenation zone is a hydrogenation reactor, and the third hydrogenation zone is a hydrogenation reactor, but the following examples do not limit the method of the present invention.

The specific implementation manner of the step-by-step diesel oil hydrogenation process based on different types of aromatic hydrocarbons shown in figure 1 is as follows: firstly, a raw oil supply unit 20 obtains a heavy fraction raw oil 1, a medium fraction raw oil 3 and a light fraction raw oil 5, wherein the heavy fraction raw oil 1 is pressurized by a first raw oil pump 2 to obtain a mixture, then the mixture is mixed with circulating hydrogen and then enters a first hydrogenation reactor 7, and passes through a hydrofining catalyst bed (or a hydrofining catalyst bed) to obtain a first hydrogenation reaction effluent 10, the medium fraction raw oil 3 is pressurized by a second raw oil pump 4, then the mixture with the first hydrogenation reaction effluent 10 enters a second hydrogenation reactor 8, and passes through the hydrofining catalyst bed (or the hydrofining catalyst bed) to obtain a second hydrogenation reaction effluent 11, the light fraction raw oil 5 is pressurized by a second raw oil pump 6 and then enters a third hydrogenation reactor 9 together with the second hydrogenation reaction effluent 11, and passes through the hydrofining catalyst bed, and the obtained third hydrogenation reaction effluent 12 enters a high-pressure separator 13, high-pressure hydrogen-rich gas 14 and liquid are obtained by separation in the high-pressure separator 13, the liquid enters a stripping/fractionating system 15, naphtha 16 and diesel oil 17 are obtained by fractionation in the stripping/fractionating system 15, and the high-pressure hydrogen-rich gas 14 is directly or after hydrogen sulfide is removed, pressurized by a recycle hydrogen compressor 18 and mixed with new hydrogen 19 to be used as recycle hydrogen.

The following examples further illustrate specific aspects of the present invention. The FHUDS-5 catalyst, the FHUDS-6 catalyst and the 3963 catalyst which are researched, developed and produced by FRIPP are used for experimental research, wherein the FHUDS-5 catalyst takes Mo-Co as a hydrogenation active center, the FHUDS-6 catalyst takes Mo-Ni as a hydrogenation active center, the 3963 catalyst is a hydrogenation modified catalyst and contains a Y-type molecular sieve.

TABLE 1 Main Properties of Diesel feedstocks

Table 2 examples process conditions and main product properties

TABLE 2

Table 3 examples process conditions and main product properties

TABLE 3

As can be seen from the examples, the diesel raw material can directly produce high-quality clean diesel products by the step-by-step diesel hydrogenation method based on different types of aromatic hydrocarbons.

The preferred embodiments of the present invention have been described above in detail, but the present invention is not limited thereto. Within the scope of the technical idea of the invention, many simple modifications can be made to the technical solution of the invention, including combinations of various technical features in any other suitable way, and these simple modifications and combinations should also be regarded as the disclosure of the invention, and all fall within the scope of the invention.

Claims

1. A step-by-step diesel oil hydrogenation method based on different types of aromatic hydrocarbon components is characterized in that: the method comprises the following steps:

2. The method of claim 1, wherein: the sulfur content in the liquid stream of the first hydrogenation reaction zone and/or the second hydrogenation reaction zone is 10 mug/g-30 mug/g.

3. The method of claim 1, wherein: the proportion of the hydrogen sulfide added value of the first reaction zone to the total hydrogen sulfide added value is 23-50%.

4. The method of claim 1, wherein: one of the following schemes is adopted:

5. The method of claim 1, wherein: the diesel raw oil is one or more of straight-run diesel, catalytic diesel, coking diesel, thermal cracking diesel, visbreaking diesel, synthetic diesel, coal tar diesel fraction, coal direct liquefaction diesel and shale oil diesel.

6. The method of claim 1, wherein: the fractionation point temperature of the light fraction and the middle fraction obtained from the raw oil of the diesel oil is 200-270 ℃.

7. The method of claim 1, wherein: the fractionation point temperature of the middle distillate and the heavy distillate obtained from the diesel raw oil is 260-320 ℃.

8. The method of claim 1, wherein: the weight ratio of the light fraction to the raw oil is 20-40%, and the weight ratio of the medium fraction to the raw oil is 25-45%.

9. The method of claim 1, wherein: the first hydrogenation reaction zone is composed of one reactor or a plurality of reactors.

10. The method of claim 1, wherein: the operating conditions of the first hydrogenation reaction zone are that the reaction pressure is 3.0MPa to 14.0MPa, and the volume space velocity of the heavy fraction diesel oil raw material oil is 0.2h^-1~10.0h^-1The average reaction temperature is 220-450 ℃, and the volume ratio of hydrogen to oil is 200: 1-2000: 1.

11. The method of claim 9, wherein: the operation conditions are that the reaction pressure is 4.0 MPa-13.0 MPa, the volume space velocity of the heavy fraction raw oil is 0.3h < -1 > -8.0 h < -1 >, the average reaction temperature is 230 ℃ to 440 ℃, and the volume ratio of hydrogen to oil is 300: 1-1800: 1.

12. The method of claim 4, wherein: the hydro-upgrading catalyst carrier used in the first hydrogenation reaction zone contains alumina and a molecular sieve; based on the weight of the catalyst, the content of the VIB group metal is 10-35 wt% calculated by oxide, the content of the VIII group metal is 3-15 wt% calculated by oxide, the content of the molecular sieve is 5-40 wt%, and the content of the alumina is 10-80 wt%; its specific surface area is 100m²/g～650m²The pore volume is 0.15mL/g to 0.50 mL/g.

13. The method of claim 3, wherein: the first hydrogenation reaction zone hydrogenation modification process conditions are as follows: the reaction pressure is 3.0MPa to 13.0MPa, the reaction temperature is 300 ℃ to 433 ℃, and the liquid hourly volume space velocity is 0.2h^-1～15.0h^-1Hydrogen oil volume ratio of 100:1 ℃2000:1。

14. The method of claim 1, wherein: the second hydrogenation reaction zone is composed of one reactor or a plurality of reactors, and a catalyst bed layer or a plurality of catalyst bed layers are arranged in the reactor.

15. The method of claim 1, wherein: the operating conditions of the second hydrogenation reaction zone are that the reaction pressure is 3.0 MPa-14.0 MPa, and the total volume space velocity of the light fraction and heavy fraction diesel raw oil is 0.2h^-1~12.0h^-1The average reaction temperature is 210-448 ℃, and the volume ratio of hydrogen to oil is 150: 1-1700: 1.

16. The method of claim 13, wherein: the operation conditions are that the reaction pressure is 4.0MPa to 13.0MPa, and the total volume airspeed of the diesel raw oil is 0.3h^-1~10.0h^-1The average reaction temperature is 220-438 ℃, and the volume ratio of hydrogen to oil is 200: 1-1300: 1.

17. The method of claim 4, wherein: the hydro-upgrading catalyst carrier used in the second hydrogenation reaction zone contains alumina and a molecular sieve; based on the weight of the catalyst, the content of the VIB group metal is 10-35 wt% calculated by oxide, the content of the VIII group metal is 3-15 wt% calculated by oxide, the content of the molecular sieve is 5-40 wt%, and the content of the alumina is 10-80 wt%; its specific surface area is 100m²/g～650m²The pore volume is 0.15mL/g to 0.50 mL/g.

18. The method of claim 3, wherein: the second hydrogenation reaction zone hydrogenation modification process conditions are as follows: the reaction pressure is 3.0MPa to 13.0MPa, the reaction temperature is 300 ℃ to 433 ℃, and the liquid hourly volume space velocity is 0.2h^-1～15.0h^-1The volume ratio of the hydrogen to the oil is 100: 1-2000: 1.

19. The method of claim 1, wherein: the average reaction temperature of the first reaction zone is 2-35 ℃ higher than the average reaction temperature of the second reaction zone.

20. The method of claim 1, wherein: the operating conditions of the third hydrogenation reaction zone are that the reaction pressure is 3.0 MPa-14.0 MPa, and the total volume space velocity of the light fraction and heavy fraction diesel raw oil is 0.2h^-1~15.0h^-1The average reaction temperature is 200-445 ℃, and the volume ratio of hydrogen to oil is 150: 1-1500: 1.

21. The method of claim 1, wherein: the average reaction temperature of the second reaction zone is 3-40 ℃ higher than the average reaction temperature of the third reaction zone.

22. The method of claim 1, wherein: and controlling the proportion of the hydrogen sulfide increase value of the second reaction zone to the total hydrogen sulfide increase value to be 26-46%.

23. The method of claim 4, wherein: the volume proportion of the extracted partial material flow in terms of gas phase accounts for 10-50 v% of the raw material gas phase at the inlet of each reaction zone.