CN111763533B

CN111763533B - Method for processing heavy distillate oil

Info

Publication number: CN111763533B
Application number: CN201910257092.4A
Authority: CN
Inventors: 胡智中; 胡志海; 赵阳; 董松涛; 吴昊; 莫昌艺
Original assignee: Sinopec Research Institute of Petroleum Processing; China Petroleum and Chemical Corp
Current assignee: Sinopec Research Institute of Petroleum Processing; China Petroleum and Chemical Corp
Priority date: 2019-04-01
Filing date: 2019-04-01
Publication date: 2021-10-08
Anticipated expiration: 2039-04-01
Also published as: CN111763533A

Abstract

The invention relates to the field of hydrocarbon oil hydrotreatment, and discloses a method for processing heavy distillate oil, which comprises the following steps: introducing the heavy distillate oil into a hydrofining reaction zone to carry out hydrofining reaction; then introducing the obtained hydrofining reaction effluent into a hydrocracking reaction zone for hydrocracking reaction; and sequentially separating and fractionating the hydrocracking reaction effluent obtained in the hydrocracking reaction zone; wherein, in the hydrocracking reaction area, the ratio of the hydrocracking activity to the cracking activity of the hydrocracking catalyst in the hydrocracking reaction area is increased according to the flowing direction of the reaction materials. The method can increase the yield of the aviation kerosene fraction on the premise of ensuring the quality of the aviation kerosene fraction and even improving the quality of the aviation kerosene fraction.

Description

Method for processing heavy distillate oil

Technical Field

The invention relates to the field of hydrocarbon oil hydrotreating, in particular to a method for processing heavy distillate oil.

Background

The hydrocracking technology has the characteristics of strong raw material adaptability, large flexibility of production operation and product scheme, good product quality and the like, can directly convert various heavy inferior feeds into high-quality jet fuel, diesel oil, chemical naphtha required by the market and tail oil used as a raw material for preparing ethylene by steam cracking, is one of the most important heavy oil deep processing technologies in modern oil refining and petrochemical industry, and is increasingly widely applied at home and abroad.

The hydrocracking comprises two processes of hydrogenation and cracking, wherein in the hydrocracking reaction process, aromatic hydrocarbon compounds are subjected to side chain scission reaction to generate paraffin with different molecular weights on one hand, and aromatic rings of polycyclic aromatic hydrocarbons are subjected to hydrogenation saturation reaction to generate monocyclic aromatic hydrocarbons which are not easy to be subjected to hydrogenation saturation on the other hand, and when the aromatic hydrocarbons in a flowing reactant are almost all monocyclic aromatic hydrocarbons, the aromatic hydrocarbon compounds are subjected to side chain scission reaction and a small amount of monocyclic aromatic hydrocarbon saturation reaction; the side chain breaking reaction of the cyclane compound generates paraffin with different molecular weight, and the ring opening reaction of the polycyclic cyclane is performed; the paraffins are then cracked into paraffins of different molecular weights.

According to the chemical reaction processes of three hydrocarbon compounds, the hydrogenation process in hydrocracking plays a main role in the saturation reaction of polycyclic aromatic hydrocarbons, mainly focuses on the initial reaction stage (the conversion per pass of distillate is more than 320 ℃) and is weakened in the subsequent reaction. Therefore, in the process, the problems of selective hydrogenation and cracking can be involved, and the aim of increasing the yield of the aviation kerosene fraction and improving the properties of the aviation kerosene and tail oil fraction can be achieved by controlling the hydrogenation activity and the cracking activity of the catalyst, namely selecting and controlling the activity of the catalyst and the corresponding reaction temperature.

CN104611019A discloses a low-energy consumption hydrocracking method for producing high-quality jet raw materials, wherein raw oil is mixed with hydrogen, and after twice heat exchange, the raw oil sequentially passes through a hydrofining reaction zone and a hydrocracking reaction zone; the hydrocracking reaction zone comprises at least two hydrocracking catalysts, wherein the upstream is filled with a catalyst I, and the downstream is filled with a catalyst II; the content of the modified Y molecular sieve in the catalyst I is 10-25 percent higher than that of the catalyst II. The method organically combines a high-temperature high-pressure countercurrent heat transfer technology with a hydrocracking catalyst grading technology, comprehensively utilizes hydrocracking reaction heat, not only fully exerts the characteristics of two different types of hydrocracking catalysts, keeps the selectivity of the catalysts, but also improves the quality of target products, and reduces engineering investment and operation energy consumption.

CN104611040A discloses a hydrocracking method, wherein heavy distillate oil is mixed with hydrogen and then enters a hydrofining reactor for hydrofining reaction; the effluent of the hydrorefining reaction directly enters a hydrocracking reactor and contacts and reacts with a graded catalyst bed layer in the hydrocracking reactor; wherein, at least two cracking catalyst beds are arranged in the hydrocracking reactor, and the hydrogenation activity of the hydrocracking catalyst is in a decreasing trend according to the flowing direction of the reaction materials; and separating and fractionating the hydrocracking reaction effluent to obtain hydrocracking products including heavy naphtha and tail oil. The method can reduce excessive hydrogenation and excessive cracking of heavy naphtha and reduce chemical hydrogen consumption while ensuring the hydrocracking effect of heavy components, thereby improving the aromatic hydrocarbon potential and selectivity of the heavy naphtha.

Disclosure of Invention

The invention aims to solve the problem of increasing the yield of the aviation kerosene fraction on the premise of ensuring the quality of the aviation kerosene fraction and even improving the quality of the aviation kerosene fraction.

In order to achieve the above object, the present invention provides a process for processing heavy distillate, the process comprising: introducing heavy distillate oil into a hydrofining reaction zone in the presence of hydrogen to perform hydrofining reaction; then introducing the hydrofining reaction effluent obtained in the hydrofining reaction zone into a hydrocracking reaction zone for hydrocracking reaction; sequentially separating and fractionating the hydrocracking reaction effluent obtained in the hydrocracking reaction zone to obtain a aviation kerosene fraction and a tail oil fraction; wherein, in the hydrocracking reaction zone, the R value of the downstream hydrocracking catalyst is more than or equal to the R value of the upstream hydrocracking catalyst according to the flowing direction of the reaction materials, and the hydrocracking reaction zone contains at least two hydrocracking catalysts with different R values, wherein the R value is the total weight of metal components of the hydrocracking catalysts calculated by oxide/the total weight of molecular sieves of the hydrocracking catalysts, and the R values of the hydrocracking catalysts are respectively and independently 2.5-6.

The present invention can increase the yield of aviation kerosene fraction while ensuring the quality of aviation kerosene fraction and even improving the quality of aviation kerosene fraction by adjusting the weight ratio of metal component (in terms of oxide) to molecular sieve of hydrocracking catalyst in the process for processing heavy distillate to be within a specific range (2.5-6, preferably 2.5-4).

Specifically, according to the flowing direction of the reaction materials, the yield of the aviation kerosene fraction can be increased on the premise of ensuring the quality of the aviation kerosene fraction and even improving the quality of the aviation kerosene fraction by controlling the ratio of the hydrogenation activity to the cracking activity of the hydrocracking catalyst in the hydrocracking reaction zone to be in an increasing trend and matching with the hydrofining reaction, the hydrocracking reaction and the like in the method for processing the heavy distillate oil.

Detailed Description

The endpoints of the ranges and any values disclosed herein are not limited to the precise range or value, and such ranges or values should be understood to encompass values close to those ranges or values. For ranges of values, between the endpoints of each of the ranges and the individual points, and between the individual points may be combined with each other to give one or more new ranges of values, and these ranges of values should be considered as specifically disclosed herein.

As previously mentioned, the present invention provides a process for processing heavy distillate.

"the R-value of the downstream hydrocracking catalyst is greater than or equal to the R-value of the upstream hydrocracking catalyst" includes, but is not limited to, the process of the present invention is carried out in a hydrocracking reaction zone containing at least 3 hydrocracking catalyst beds, wherein, in terms of the direction of flow of the reaction mass, the R-value of the first hydrocracking catalyst bed is equal to the R-value of the second hydrocracking catalyst bed, and the R-value of the second hydrocracking catalyst bed is less than the R-value of the third hydrocracking catalyst bed; or the R value of the first hydrocracking catalyst bed layer is smaller than that of the second hydrocracking catalyst bed layer, and the R value of the second hydrocracking catalyst bed layer is equal to that of the third hydrocracking catalyst bed layer. The process of the present invention is similarly explained when it is carried out in a hydrocracking reaction zone containing at least 4 beds of hydrocracking catalyst.

The method for processing the heavy distillate oil can better perform selective hydrogenation and cracking, and can improve the properties of the aviation kerosene fraction and the tail oil fraction while increasing the yield of the aviation kerosene fraction.

Preferably, in said hydrocracking reaction zone, each of said hydrocracking catalysts independently has an R value of from 2.5 to 4. The inventors of the present invention have found that the quality of the aviation kerosene fraction obtained by the process of the present invention is higher when the R-values of the respective hydrocracking catalysts are each independently 2.5 to 4.

Preferably, in the present invention, the hydrofinishing reaction effluent is directly introduced into the hydrocracking reaction zone without further separation operation for hydrocracking reaction.

According to a preferred embodiment, at least two hydrocracking catalyst beds are arranged in the hydrocracking reaction zone, and the R value of the hydrocracking catalyst in the downstream hydrocracking catalyst bed is larger than that of the hydrocracking catalyst in the upstream hydrocracking catalyst bed according to the flowing direction of the reaction materials. In this preferred embodiment, the hydrocracking catalysts in the same hydrocracking catalyst bed are of the same type.

That is, in the present invention, the hydrogenation activity and the cracking activity of the hydrocracking catalyst are represented by parts by weight of a metal component in terms of oxide contained in the hydrocracking catalyst and parts by weight of a molecular sieve component contained therein, respectively.

According to a preferred embodiment, in the hydrocracking reaction zone, the total weight fraction of metal components in terms of oxides in the downstream hydrocracking catalyst is increased, unchanged or decreased relative to the upstream hydrocracking catalyst in terms of the flow direction of the reaction mass.

According to another preferred embodiment, in the hydrocracking reaction zone, the total weight fraction of the molecular sieve in the downstream hydrocracking catalyst is increased, unchanged or decreased relative to the upstream hydrocracking catalyst, in terms of the flow direction of the reaction mass.

That is, in the present invention, the purpose of adjusting the R value of the hydrocracking catalyst is achieved by adjusting the total weight part of the metal component in the hydrocracking catalyst and/or the total weight part of the molecular sieve.

Preferably, the content of the metal component in the hydrocracking catalyst calculated by oxide is 5 wt% -60 wt%, and the content of the molecular sieve is 5 wt% -40 wt% based on the total weight of the hydrocracking catalyst; more preferably, the content of the metal component in terms of oxide in the hydrocracking catalyst is 10 wt% to 50 wt%, and the content of the molecular sieve is 5 wt% to 25 wt%, based on the total weight of the hydrocracking catalyst.

According to a preferred embodiment, the molecular sieve has a specific surface area of 600m²/g-1000m²The total pore volume is 0.33mL/g to 0.50mL/g, the relative crystallinity is 80 percent to 140 percent, the unit cell parameter is 2.427 to 2.480, and the infrared acid amount is 0.3 to 1.8 mmol/g.

In the present invention, the molecular sieve is more preferably a Y-type molecular sieve and/or a β -type molecular sieve.

Preferably, the metal component in the hydrocracking catalyst includes a W element, and the metal component further includes at least one of a Ni element and a Co element.

Preferably, the content of the element W in terms of oxide in the hydrocracking catalyst is 10 wt% -35 wt%, and the total content of the element Ni and the element Co is 5 wt% -20 wt%, based on the total weight of the hydrocracking catalyst.

Preferably, in the hydrocracking reaction zone, the volume ratio of any two adjacent hydrocracking catalyst beds is 1: (0.2-5), more preferably 1: (0.5-2).

Preferably, the hydrocracking catalyst of the present invention contains alumina and/or amorphous silica alumina as a carrier.

In the process of the present invention, the hydrorefining catalyst and hydrocracking agent to be used may be those commercially available or may be those prepared according to conventional knowledge in the art.

The hydrofining catalyst used in the invention can adopt a conventional hydrocracking pretreatment catalyst. Preferably, the hydrorefining catalyst contains a hydrogenation active component and a carrier, the hydrogenation active component is at least one of VIB group metal elements and VIII group metal elements, and the carrier contains alumina or silicon-containing alumina.

More preferably, in the hydrorefining catalyst, the group VIB metal element is Mo and/or W, and the group VIII metal element is Co and/or Ni.

Preferably, the content of the VIB group metal element in terms of oxide in the hydrofining catalyst is 5 wt% -35 wt%, and the content of the VIII group metal element in terms of oxide in the hydrofining catalyst is 1 wt% -20 wt%.

Preferably, the reaction conditions in the hydrofinishing reaction zone include: the reaction temperature is 300-450 ℃, and is more preferably 330-370 ℃; the reaction pressure is 6.0MPa-17.0MPa, more preferably 8.0MPa-14.0 MPa; the liquid hourly space velocity is 0.1h^-1-5.0h^-1More preferably 0.5h^-1-1.5h^-1(ii) a The hydrogen-oil volume ratio in the standard state is 300-2000, and more preferably 600-1000.

Preferably, the reaction conditions in the hydrocracking reaction zone include: the reaction temperature is 310-470 ℃, and more preferably 350-400 ℃; the reaction pressure is 5.0MPa-20.0MPa, more preferably 10.0MPa-16.0 MPa; the liquid hourly space velocity is 0.2h^-1-6.0h^-1More preferably 0.7h^-1-2.5h^-1(ii) a The hydrogen-oil volume ratio at the standard state is 100-.

According to a particularly preferred embodiment, in the hydrocracking reaction zone, the temperature difference between any two adjacent hydrocracking catalyst beds is independently from 5 ℃ to 30 ℃, more preferably from 10 ℃ to 15 ℃, in terms of the flow direction of the reaction material, and the temperature of the downstream hydrocracking catalyst bed is higher than the temperature of the upstream hydrocracking catalyst bed. The inventors of the present invention have found that with this preferred embodiment, particularly when the temperature difference between any two adjacent hydrocracking catalyst beds is independently 10 ℃ to 15 ℃, the yield of the aviation kerosene fraction obtained by the process of the present invention is higher and the quality is better.

More preferably, in the present invention, the reaction temperature of the hydrocracking catalyst is gradually increased from top to bottom as the hydrocracking catalyst hydrogenation activity/cracking activity increases, according to the flow direction of the reaction material.

Preferably, the reaction conditions in the hydrocracking reaction zone are controlled such that the conversion per pass of the tail oil fraction during the hydrocracking reaction is in the range of 50% to 80%, more preferably in the range of 58% to 69%.

Preferably, the terminal boiling point of the aviation kerosene fraction is 260 ℃ to 295 ℃ and the initial boiling point of the tail oil fraction is >320 ℃. Particularly preferably, the distillation range of the aviation kerosene fraction is 150-280 ℃, and the distillation range of the tail oil fraction is more than 320 ℃.

Compared with the prior art, the method for processing the heavy distillate oil has the following specific advantages:

in a hydrocracking reaction zone, the catalyst hydrogenation activity/cracking activity (total metal oxide amount/molecular sieve content) is increased from top to bottom along the material flow direction, and the R values of various hydrocracking catalysts are controlled to be respectively and independently 2.5-6, more preferably respectively and independently 2.5-4, and further preferably, the obvious effects on the increase of the aviation kerosene smoke point and the reduction of the BMCI value of tail oil are achieved by properly increasing the reaction temperature (preferably, the temperature difference is 5-30 ℃ and more preferably 10-15 ℃) of the reaction section; the method can improve the properties of the aviation kerosene fraction and the tail oil fraction on the premise of improving the aviation kerosene yield.

In the invention, in the hydrocracking reaction zone, because the hydrogenation reaction is a strong exothermic reaction and the exothermic effect is greater than the endothermic effect of hydrocarbon molecule cracking, the temperature rise phenomenon occurs between catalyst beds in the reaction zone, and in order to maintain the stability of the selectivity of the hydrocracking catalyst industrially, cold hydrogen is introduced between the catalyst beds to cool the catalyst beds, thereby increasing the energy consumption of the hydrocracking device. Under the optimal condition, the reaction temperature of a downstream bed layer is properly increased in the invention, so that the injection of cold hydrogen in the hydrocracking device can be effectively reduced, and the energy consumption of the device is reduced.

The present invention is not particularly limited with respect to the source of the heavy distillate, and may be various heavy distillates conventionally used in the art. Preferably, the heavy distillate oil is selected from one or more of vacuum wax oil, deasphalted oil, catalytic cracking light cycle oil and coking wax oil; further preferably, the initial boiling point of the heavy distillate oil is 180-250 ℃, and the final boiling point of the heavy distillate oil is 500-600 ℃.

In the method of the present invention, various kinds of hydrogenation protection catalysts conventionally used in the art may be further packed upstream of the hydrofinishing reaction zone and/or downstream of the hydrocracking reaction zone, and the packing volume of the hydrogenation protection catalyst is not particularly limited, and those skilled in the art can determine the loading volume according to actual needs by combining conventional operation means in the art.

In the present invention, the conversion per pass is (tail fraction in the feed oil-tail fraction in the product oil)/tail fraction in the feed oil × 100%.

The present invention will be described in detail below by way of examples. In the following examples, the raw materials used are all commercially available ones unless otherwise specified.

The catalyst loading in the hydrofinishing reaction zone in the following examples was the same and the total catalyst loading in the hydrocracking reaction zone was the same.

The hydrorefining catalyst used below was a catalyst produced by Changjingtie of petrochemical Co., Ltd., China, having model number RN-410.

The hydrocracking catalysts used below were prepared by all of the conventional impregnation methods, and the compositions and main properties of the obtained hydrocracking catalysts are listed in table 1. Wherein the specific surface area of the Y-type molecular sieve is 720m²The total pore volume is 0.39mL/g, the relative crystallinity is 98 percent, the unit cell parameter is 2.441, and the infrared acid amount is 0.6 mmol/g; the specific surface area of the beta type molecular sieve is 760m²Per g, total pore volume of 041mL/g, 95% relative crystallinity, 2.452 unit cell parameter, 0.55mmol/g infrared acid. Both molecular sieves are from Changjingtie, Inc., petrochemical Co.

The impregnation method comprises the following specific operation processes:

weighing molecular sieve, carrier and binder (such as sesbania powder), grinding and mixing them uniformly, adding dilute nitric acid solution, stirring to obtain jelly with uniform humidity, extruding into strips on a strip extruder, and sequentially drying and roasting to obtain the hydrocracking carrier. Then, the active component is impregnated on the hydrocracking carrier by adopting a saturated impregnation method. Mixing the prepared solution containing the active metal component with the hydrocracking carrier, stirring the mixture at uniform humidity, standing the mixture at normal temperature for a certain time, drying the catalyst at 120 ℃ for 5 hours, heating the mixture to 450 ℃ at the speed of 3.0 ℃/min, and roasting the mixture at constant temperature for 4 hours to prepare the hydrocracking catalyst, wherein the catalyst A is numbered.

Other catalyst preparation methods are essentially the same as the above operations, differing only in changing the molecular sieve content or type and the amount of supported metal.

The properties of the heavy distillate used below are listed in table 2.

TABLE 1

Wherein, the silicon-containing alumina in the catalyst E contains silicon oxide in molar ratio: 3:1 of alumina

TABLE 2

Item	Heavy distillate oil	Hydrorefining the reaction effluent
			Density, g/cm³	0.9075	0.8731
Distillation range, deg.C	233-539	133-515
			C，wt％	85.35	86.75
H，wt％	12.27	13.24
			S，wt％	2.47	0.116
N，wt％	0.071	0.001
			BMCI value	46.7	33.0

Example 1

Introducing hydrogen and heavy distillate oil into a hydrofining reactor containing a hydrofining catalyst to carry out hydrofining reaction, and controlling the content of N in the effluent of the hydrofining reaction to be 10 ppm; then, the hydrofining reaction effluent (with properties shown in table 2) obtained in the hydrofining reactor is directly introduced into a hydrocracking reactor for hydrocracking reaction; sequentially separating and fractionating the hydrocracking reaction effluent obtained in the hydrocracking reactor to obtain a aviation kerosene fraction with the distillation range of 150-280 ℃ and a tail oil fraction with the distillation range of more than 320 ℃; the hydrocracking reactor contains 2 hydrocracking catalyst beds, and different hydrocracking catalysts are filled in an upstream hydrocracking catalyst bed and a downstream hydrocracking catalyst bed according to the flowing direction of reaction materials, and are specifically shown in table 4.

The reaction conditions in the hydrofinishing reactor of this example are listed in table 3, and the reaction conditions in the hydrocracking reactor are listed in table 4, and the main properties of the obtained product are listed in table 5.

Example 2

This example was carried out by a similar process flow to that of example 1, except that the kind of catalyst, the packed volume ratio, the reaction temperature of different beds, and the like in the hydrocracking reaction zone in this example were different from those of example 1, specifically, the reaction conditions in the hydrocracking reactor of this example are shown in table 4, and the main properties of the obtained product are shown in table 5.

Example 3

This example was carried out using a similar process scheme to that of example 1, except that the reaction temperature of the different beds in this example was different from that of example 1, specifically, the reaction conditions in the hydrocracking reactor of this example are shown in table 4, and the main properties of the obtained product are shown in table 5.

Example 4

This example was carried out by a process similar to that of example 1, except that 3 hydrocracking catalyst beds were disposed in the hydrocracking reaction zone in this example, the type of catalyst, the packing volume ratio, the reaction temperatures of the different beds, and the like in the hydrocracking reaction zone in this example were different from those in example 1, specifically, the reaction conditions in the hydrocracking reactor in this example are shown in table 4, and the main properties of the obtained product are shown in table 5.

Example 5

This example was carried out using a similar process scheme to that of example 1, except that the type of catalyst in the hydrocracking reaction zone in this example was different from that of example 1. Specifically, the reaction conditions in the hydrocracking reactor of this example are listed in table 4, and the main properties of the obtained product are listed in table 5.

Comparative example 1

This comparative example was conducted using a similar process flow to that of example 1, except that the same hydrocracking catalyst was packed in 2 hydrocracking catalyst beds disposed in the hydrocracking reaction zone in this comparative example, specifically, the reaction conditions in the hydrocracking reactor of this comparative example are shown in table 4, and the main properties of the obtained product are shown in table 5.

Comparative example 2

This comparative example was conducted using a similar process flow to that of example 4, except that the same hydrocracking catalyst was packed in 3 hydrocracking catalyst beds disposed in the hydrocracking reaction zone in this comparative example, specifically, the reaction conditions in the hydrocracking reactor of this comparative example are shown in table 4, and the main properties of the obtained product are shown in table 5.

Comparative example 3

This comparative example was conducted by a similar process flow to that of example 3, except that the type of hydrocracking catalyst packed in the 2 hydrocracking catalyst beds disposed in the hydrocracking reaction zone in this comparative example was different, specifically, the reaction conditions in the hydrocracking reactor of this comparative example are shown in table 4, and the main properties of the obtained product are shown in table 5.

Comparative example 4

This comparative example was conducted by a similar process flow to that of example 1, except that the types of hydrocracking catalysts packed in the 2 hydrocracking catalyst beds disposed in the hydrocracking reaction zone in this comparative example were different, specifically, the reaction conditions in the hydrocracking reactor of this comparative example are listed in table 4, and the main properties of the obtained product are listed in table 5.

TABLE 3

Reaction conditions in the hydrofining reaction zone
		Temperature, C	365
Volumetric space velocity h^-1	1.0
		Hydrogen to oil ratio at standard state volume	800
Reaction pressure, MPa	14.0

TABLE 4

TABLE 5

As can be seen from the comparison results, the activity of the hydrocracking catalyst was independently adjusted under the same >320 ℃ fraction per pass conversion (60%) (example 1), the yield of the aviation kerosene fraction (150 ℃ to 280 ℃) was improved by 5.12% compared with that of the single catalyst and the isothermal condition (comparative example 1), and the yield of the aviation kerosene fraction was improved by about 7.72% compared with that of the comparative example by controlling the R value of the hydrocracking catalyst and the reaction temperature of the corresponding catalyst (example 3).

In addition, the results of the invention show that the method provided by the invention has the advantages of high aviation soot smoke point and low BMCI value of tail oil. That is, the method of the invention can obtain higher-quality aviation kerosene products and ethylene raw materials prepared by steam cracking.

In conclusion, the method can increase the yield of the aviation kerosene fraction on the premise of ensuring the quality of the aviation kerosene fraction and even improving the quality of the aviation kerosene fraction.

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 process for processing heavy distillate, the process comprising: introducing heavy distillate oil into a hydrofining reaction zone in the presence of hydrogen to perform hydrofining reaction; then introducing the hydrofining reaction effluent obtained in the hydrofining reaction zone into a hydrocracking reaction zone for hydrocracking reaction; sequentially separating and fractionating the hydrocracking reaction effluent obtained in the hydrocracking reaction zone to obtain a aviation kerosene fraction and a tail oil fraction; wherein, in the hydrocracking reaction zone, the R value of the downstream hydrocracking catalyst is more than or equal to the R value of the upstream hydrocracking catalyst according to the flowing direction of the reaction materials, and the hydrocracking reaction zone contains at least two hydrocracking catalysts with different R values, wherein the R value is the total weight of metal components of the hydrocracking catalysts calculated by oxide/the total weight of molecular sieves of the hydrocracking catalysts, and the R values of the hydrocracking catalysts are respectively and independently 2.5-6.

2. The process of claim 1, wherein each of said hydrocracking catalysts has an R value of from 2.5 to 4 independently in said hydrocracking reaction zone.

3. The process of claim 1, wherein at least two hydrocracking catalyst beds are arranged in the hydrocracking reaction zone, and the R value of the hydrocracking catalyst in the downstream hydrocracking catalyst bed is larger than that of the hydrocracking catalyst in the upstream hydrocracking catalyst bed according to the flow direction of the reaction materials.

4. The process as claimed in any one of claims 1 to 3, wherein in the hydrocracking reaction zone, the total weight fraction of metal components in terms of oxides in the downstream hydrocracking catalyst is increased, unchanged or decreased relative to the upstream hydrocracking catalyst in terms of the flow direction of the reaction mass.

5. The process of any one of claims 1 to 3, wherein the total weight fraction of molecular sieve in the downstream hydrocracking catalyst relative to the upstream hydrocracking catalyst is increased, unchanged or decreased in the hydrocracking reaction zone in terms of the direction of flow of the reaction mass.

6. The process according to any one of claims 1 to 3, wherein the hydrocracking catalyst contains a metal component in an amount of 5 wt% to 60 wt% in terms of oxide and a molecular sieve in an amount of 5 wt% to 40 wt%, based on the total weight of the hydrocracking catalyst.

7. The process of claim 6, wherein the hydrocracking catalyst contains a metal component in an oxide amount of 10 wt% to 50 wt% and a molecular sieve amount of 5 wt% to 25 wt%, based on the total weight of the hydrocracking catalyst.

8. The method of claim 6, wherein the molecular sieve has a specific surface area of 600m²/g-1000m²Per g, total pore volume 0.33mL/g to 0.50mL/g, relative crystallinity of 80% to 140%, unit cell parameters of 2.427 to 2.480, infrared acid amount of 0.3 to 1.8 mmol/g.

9. The method of claim 6, wherein the metal component in the hydrocracking catalyst comprises a W element, and the metal component further comprises at least one of a Ni element and a Co element.

10. The method according to claim 9, wherein the hydrocracking catalyst contains W element in an amount of 10 wt% to 35 wt% in terms of oxide and Ni element and Co element in an amount of 5 wt% to 20 wt% in total, based on the total weight of the hydrocracking catalyst.

11. The process of any one of claims 1 to 3, wherein in the hydrocracking reaction zone, the volume ratio of any two adjacent hydrocracking catalyst beds is each independently 1: (0.2-5).

12. The process of claim 11, wherein in the hydrocracking reaction zone, the volume ratio of any two adjacent hydrocracking catalyst beds is each independently 1: (0.5-2).

13. The process of any one of claims 1-3, wherein the hydrofinishing reaction zone is filled with a hydrofinishing catalyst, the hydrofinishing catalyst comprises a hydrofinishing active component and a carrier, the hydrofinishing active component is at least one of a group VIB metal element and a group VIII metal element, and the carrier comprises alumina or silicon-containing alumina.

14. The process according to claim 13, wherein in the hydrofinishing catalyst, the group VIB metal elements are Mo and/or W and the group VIII metal elements are Co and/or Ni.

15. The process according to claim 13, wherein the hydrofinishing catalyst comprises 5 to 35 wt% of the group VIB metal element calculated as oxides and 1 to 20 wt% of the group VIII metal element calculated as oxides, based on the total weight of the hydrofinishing catalyst.

16. The process of any of claims 1-3, wherein the reaction conditions in the hydrofinishing reaction zone comprise: the reaction temperature is 300-450 ℃; the reaction pressure is 6.0MPa-17.0 MPa; the liquid hourly space velocity is 0.1h^-1-5.0h^-1(ii) a The hydrogen-oil volume ratio under the standard state is 300-2000.

17. The process of claim 16, wherein the reaction conditions in the hydrofinishing reaction zone comprise: the reaction temperature is 330-370 ℃; the reaction pressure is 8.0MPa-14.0 MPa; the liquid hourly space velocity is 0.5h^-1-1.5h^-1(ii) a The volume ratio of hydrogen to oil in the standard state is 600-1000.

18. The process of any one of claims 1-3, wherein the reaction conditions in the hydrocracking reaction zone comprise: the reaction temperature is 310-470 ℃; the reaction pressure is 5.0MPa-20.0 MPa; the liquid hourly space velocity is 0.2h^-1-6.0h^-1(ii) a The hydrogen-oil volume ratio under the standard state is 100-.

19. The process of claim 18, wherein the reaction conditions in the hydrocracking reaction zone comprise: the reaction temperature is 350-400 ℃; the reaction pressure is 10.0MPa-16.0 MPa; the liquid hourly space velocity is 0.7h^-1-2.5h^-1(ii) a The hydrogen-oil volume ratio under the standard state is 700-1300.

20. The process according to any one of claims 1 to 3, wherein in the hydrocracking reaction zone, the temperature difference between any two adjacent hydrocracking catalyst beds is each independently 5 ℃ to 30 ℃ in terms of the flow direction of the reaction material, and the temperature of the downstream hydrocracking catalyst bed is higher than the temperature of the upstream hydrocracking catalyst bed.

21. The process according to any one of claims 1 to 3, wherein in the hydrocracking reaction zone, the temperature difference between any two adjacent hydrocracking catalyst beds is each independently 10 ℃ to 15 ℃ in terms of the flow direction of the reaction material, and the temperature of the downstream hydrocracking catalyst bed is higher than the temperature of the upstream hydrocracking catalyst bed.

22. The process of any one of claims 1 to 3, wherein the reaction conditions in the hydrocracking reaction zone are controlled such that the conversion per pass of the tail oil fraction during the hydrocracking reaction is in the range of 50% to 80%.

23. The process of any one of claims 1 to 3, wherein the reaction conditions in the hydrocracking reaction zone are controlled such that the conversion per pass of the tail oil fraction during the hydrocracking reaction is in the range of from 58% to 69%.

24. A process according to any one of claims 1 to 3, wherein the terminal boiling point of the aviation kerosene fraction is 260 ℃ to 295 ℃ and the initial boiling point of the tail oil fraction is >320 ℃.

25. The method of any one of claims 1-3, wherein the heavy distillate oil is selected from one or more of vacuum wax oil, deasphalted oil, catalytic cracking light cycle oil and coking wax oil.

26. The process of any of claims 1-3, wherein the initial boiling point of the heavy distillate is from 180 ℃ to 250 ℃ and the final boiling point of the heavy distillate is from 500 ℃ to 600 ℃.