Hydrocracking method for preparing heavy naphtha

Abstract

The present invention discloses a hydrocracking method for preparing heavy naphtha. The method includes a hydropre-refining reaction zone and a hydrocracking reaction zone, and the process includes: diesel feedstock, hydrogen and circulating oil are mixed and fed from the top of the hydropre-refining reaction zone, coking gasoline feedstock is fed from the middle and lower part of the hydropretreatment reaction zone, and the hydropre-refining reaction is carried out. The effluent of the hydropre-refining reaction zone is mixed with at least part of liquefied gas and at least part of light naphtha and then enters the hydrocracking reaction zone for reaction. The effluent of the hydrocracking reaction zone is separated and fractionated to obtain dry gas, liquefied gas, light naphtha, heavy naphtha and diesel, wherein the diesel is circulated back to the hydropre-refining reaction zone as the circulating oil, and at least part of the liquefied gas and at least part of the light naphtha are circulated back to the hydrocracking reaction zone. The method of the present invention can improve the quality of heavy naphtha and realize the maximum production of heavy naphtha by single-stage full-cycle processing of hydrocarbon feedstock.

Description

Technical Field

[0001] The present invention relates to a hydrocracking method, in particular to a hydrocracking method for producing heavy naphtha. Background Art

[0002] From 2010 to 2014, China's domestic gasoline consumption demand was strong, and the apparent consumption achieved rapid growth, with a five-year compound growth rate of 9.46%. Since 2015, the growth rate of gasoline apparent consumption has slowed down. Affected by the increase in the purchase of small-displacement cars and the rapid development of new energy vehicles, China's apparent gasoline consumption in 2016 was 120 million tons, a year-on-year increase of 3.8%, a decrease of 5.7 percentage points from 2015; affected by the decline in domestic demand growth, the growth rate of gasoline production slowed down accordingly, with production in 2016 reaching 129 million tons, a year-on-year increase of 6.8%, a decrease of 2.9 percentage points from 2015; exports were 9.6929 million tons, a significant year-on-year increase of 66.8%; imports were 207,700 tons, a year-on-year increase of 22.1%; net exports reached 9.4852 million tons, a year-on-year increase of 68.1%. Although on the surface, the growth rate has slightly decreased, its trend is still increasing every year, so there is still great potential for increasing gasoline production in China in the short term.

[0003] Hydrocracking process is one of the two powerful tools for adjusting product structure, and hydrocracking technology has been increasingly valued by people because of its more environmentally friendly product quality. Therefore, making good use of heavy naphtha, one of the raw materials for increasing gasoline production in the largest amount by hydrocracking technology, is an important direction for technology research and development.

[0004] To increase the production of heavy naphtha through hydrocracking technology, we need to start from multiple aspects. The core catalyst selection, process optimization, raw material characteristics, parameter settings, etc. are all within the scope of optimization. For increasing the production of heavy naphtha, the processed raw materials are mostly straight-run wax oil. How to optimize the processing to achieve the maximum increase in heavy naphtha production is a difficult problem that needs to be overcome.

[0005] Straight-run diesel accounts for a relatively high proportion of crude oil cutting products. With the changes in market demand, it has become a "chicken rib" in the adjustment of refining product structure. How to use hydrocracking technology to process straight-run diesel, especially with the help of inferior diesel, with the goal of maximizing the production of high-quality gasoline, is worthy of in-depth study.

[0006] The coking process is one of the representative technologies of the thermal processing process of oil refining. It is a thermal processing process that uses residual oil as raw material and performs deep thermal cracking reaction at high temperature. The products of the reaction are gas, gasoline, diesel, wax oil and coke. Among them, the coking gasoline has a high content of unsaturated hydrocarbons and contains a large amount of non-hydrocarbon compounds such as sulfur and nitrogen, so its stability is very poor and must be subsequently processed before it can be used as engine fuel.

[0007] CN201811279942.2 discloses a method for producing heavy naphtha and jet fuel, in which the raw oil is mixed with hydrogen and then passed through a hydrogenation pre-refining reaction zone and a hydrocracking reaction zone in sequence, and the obtained reaction product is separated to obtain heavy naphtha and jet fuel, and the hydrocracking reaction zone is filled with two hydrocracking catalysts from top to bottom, and the total catalyst volume of the hydrocracking reaction zone is 100%, the volume fraction of the hydrocracking catalyst I is 50% to 85%, and the volume fraction of the hydrocracking catalyst II is 15% to 50%. The method provided by the present invention can take into account the activity and stability of the hydrocracking catalyst, and produce heavy naphtha and jet fuel components in the largest amount at a high conversion rate.

[0008] CN200610008415.9 discloses a hydrocracking method. The method comprises the following steps: pre-refining heavy distillate oil and inferior catalytic cracking diesel oil by hydrogenation, mixing the obtained streams, and entering a hydrocracking reactor. The obtained cracking products are separated into gas and liquid, and the liquid phase is fractionated to obtain light naphtha, heavy naphtha, jet fuel, diesel oil and tail oil. The inferior catalytic cracking diesel oil used has a density of more than 0.9 g / ml, an aromatic content of more than 60 wt%, and a cetane number of less than 30. The method makes full use of the characteristics of the hydrocracking process, and uses heavy distillate oil and inferior catalytic cracking diesel oil as raw materials to produce high aromatic potential heavy naphtha and high-quality tail oil, which can be used as high-quality catalytic reforming raw materials and steam cracking raw materials for ethylene production.

[0009] In the above methods, either different catalysts are used in the cracking stage for grading, or the raw material characteristics are used for staged processing. However, further research is needed on how to produce heavy naphtha. Summary of the invention

[0010] When producing heavy naphtha by conventional single-stage hydrocracking method, it is usually considered to circulate the fraction heavier than heavy naphtha to the hydrocracking reactor for further cracking to increase the production of heavy naphtha, which is generally called a single-stage full-cycle hydrocracking process. The inventors of the present invention have found through extensive research that on the basis of the single-stage full-cycle hydrocracking process, part of the liquefied gas and light naphtha fractions are used as the circulating oil of the hydrocracking reactor and the circulating amount is controlled. The circulating liquefied gas and light naphtha fractions can be completely dissolved in the circulating oil, and will not cause non-ideal processes such as gasification. The reaction balance of the whole fraction can be effectively controlled. At the same time, coking gasoline is introduced into the rear section of the hydrogenation pretreatment reaction zone to change the equilibrium point of aromatic saturation, which can achieve the purpose of increasing the production of heavy naphtha and improving the quality of heavy naphtha, thereby realizing the present invention.

[0011] In view of the problems existing in the prior art, the method of the present invention provides a hydrocracking method for preparing heavy naphtha, which can improve the quality of heavy naphtha and realize the maximum production of heavy naphtha by single-stage full-cycle processing of hydrocarbon raw materials.

[0012] The present invention provides a hydrocracking method for preparing heavy naphtha. The method comprises a hydropre-refining reaction zone and a hydrocracking reaction zone. The process comprises: diesel feedstock, hydrogen and circulating oil are mixed and fed from the top of the hydropre-refining reaction zone, coking gasoline feedstock is fed from the middle and lower part of the hydropretreatment reaction zone, and the effluent of the hydropre-refining reaction zone is mixed with at least part of liquefied gas and at least part of light naphtha and then enters the hydrocracking reaction zone for reaction. The effluent of the hydrocracking reaction zone is separated and fractionated to obtain dry gas, liquefied gas, light naphtha, heavy naphtha and diesel, wherein the diesel is circulated back to the hydropre-refining reaction zone as the circulating oil, and at least part of the liquefied gas and at least part of the light naphtha are circulated back to the hydrocracking reaction zone.

[0013] In the above hydrocracking method, the tail oil is a diesel fraction heavier than heavy naphtha, generally a fraction with an initial distillation point higher than 177°C.

[0014] In the hydrocracking method, the hydrocracking reaction zone is provided with a plurality of hydrocracking catalyst beds, preferably 3 or more, and more preferably 4 to 6. The beds may be divided by a cold hydrogen box.

[0015] In the above hydrocracking method, hydrogen is separated at the top of the separation system and can be circulated back to the hydrogenation pre-refining reaction zone and the hydrocracking reaction zone as circulating hydrogen.

[0016] In the above hydrocracking method, the middle and lower part of the hydrogenation pre-refining reaction zone refers to: the height of the middle and lower feeding position from the bottom of the last catalyst bed in the hydrogenation pre-refining reaction zone accounts for 15% to 50% of the total height of the reaction zone, preferably 20% to 45%. The total height of the reaction zone is calculated by the total height of the catalyst loaded in the hydrogenation pre-refining reaction zone.

[0017] In the above hydrocracking method, a refined catalyst is loaded in the hydrogenation pre-refining reaction zone, and the refined catalyst includes a carrier and a hydrogenation metal component. The hydrogenation metal includes a Group VIB metal and / or a Group VIII metal. The carrier is an inorganic refractory oxide, which can be selected from one or more of alumina, amorphous silica-alumina, silicon dioxide, titanium oxide, etc. The Group VIB metal is preferably W and / or Mo, and the Group VIII metal is preferably Ni and / or Co. Based on the weight of the catalyst, the content of the Group VIB metal component in terms of metal oxide is 9wt% to 35wt%, preferably 10wt% to 30wt%; the content of the Group VIII metal in terms of metal oxide is 0.5wt% to 7wt%, preferably 1.1wt% to 6wt%. The refined catalyst can be various commercial catalysts, and the required hydrogenation pre-refining catalyst can also be prepared according to the common sense in the art as needed. The main commercial catalysts include: FF-56, FF-66 developed by Dalian Petrochemical Research Institute and HC-T, HC-K, HC-P developed by UOP.

[0018] In the above hydrocracking method, a hydrocracking catalyst is loaded in the hydrocracking reaction zone, and the hydrocracking catalyst includes a carrier and a hydrogenation metal. The carrier contains a molecular sieve, and the molecular sieve includes but is not limited to a Y-type molecular sieve, such as a USY-type molecular sieve; the hydrogenation metal is selected from one or more of the metals of Group VI, Group VII, Group VIII, metal oxides or metal sulfides, preferably one or more of iron, chromium, molybdenum, tungsten, cobalt, nickel, or their sulfides or oxides. Based on the weight of the catalyst, the content of the hydrogenation metal in terms of oxide is 12wt% to 33wt%, preferably 14wt% to 32wt%, and the content of the molecular sieve is 1wt% to 30wt%. The hydrocracking catalyst can be selected from an existing commercial hydrocracking catalyst, or a specific hydrocracking catalyst can be prepared as needed. The commercial hydrocracking catalysts that can be used mainly include: FC-26, FC-50 / FC-60, etc. developed by Dalian Petrochemical Research Institute. The hydrocracking catalyst required for the process of the present invention can also be prepared according to common sense in the art as needed.

[0019] In the above-mentioned hydrocracking method, the diesel feedstock can be one or more of straight-run diesel, fluidized bed diesel, catalytic diesel, coking diesel, etc., preferably a diesel feedstock containing straight-run diesel. Furthermore, the straight-run diesel accounts for more than 50wt% of the total mass of the diesel feedstock, preferably 51wt% to 100wt%.

[0020] In the above hydrocracking method, the nitrogen content in the straight-run diesel is below 0.06wt%, preferably below 0.03wt%, further 0.001wt% to 0.029wt%, the sulfur content is not limited, the initial boiling point is generally 155 to 255°C, the final boiling point is generally 315 to 385°C, and the aromatic content is 5wt% to 35wt%, preferably 6wt% to 34wt%. The straight-run diesel can be selected from the atmospheric diesel fractions (normal second line, normal third line) obtained by processing various types of crude oil, such as one or a mixture of various atmospheric distillate diesels (such as normal second and third lines) obtained by processing Iranian crude oil or Saudi crude oil or Daqing crude oil or Karamay crude oil, such as Saudi Arabian normal second and third line diesel, Iranian normal second and third line diesel, Karamay normal second and third line diesel, Daqing normal second and third line diesel, any straight-run diesel fraction that meets the conventional requirements of the hydrocracking unit raw materials is within the scope of application of the present invention.

[0021] In the above hydrocracking method, the feed mass ratio of the diesel feedstock to the coker gasoline feedstock is 8:1 to 2:1, preferably 6:1 to 3:1.

[0022] In the above hydrocracking method, the coker gasoline generally has an initial boiling point of 22 to 45°C, preferably 25 to 40°C, a final boiling point of 160 to 205°C, preferably 170 to 200°C, an olefin content of generally 30 wt%, preferably 35 wt%, and further 36 wt% to 70 wt%; the density (20°C) of the raw material is generally 0.76 g·cm -3 Below, preferably 0.75 g·cm -3 Below, further 0.649 g·cm -3 ~0.749g·cm -3 The coker gasoline can be one of the fractionation products obtained by processing any suitable oil in a delayed coking unit, for example, it can be coker gasoline obtained by processing heavy components of Middle East crude oil, specifically, it can be coker gasoline obtained by processing Kuwait crude oil, Saudi crude oil, etc., or it can be selected from coker gasoline obtained by processing heavy components of Chinese crude oil, specifically, it can be coker gasoline obtained by processing Karamay crude oil, Daqing crude oil, etc.

[0023] In the above hydrocracking method, for the diesel feedstock at the inlet of the reactor, the main process conditions of the hydrogenation pre-refining reaction zone are as follows: average reaction temperature 260°C to 390°C, inlet hydrogen-to-oil volume ratio 350:1 to 900:1, operating pressure 6.0MPa to 10.0MPa, liquid hourly volume space velocity 0.3h-1 to 5.0h-1; preferably, the average reaction temperature is 270°C to 370°C, the inlet hydrogen-to-oil volume ratio is 400:1 to 850:1, the operating pressure is 6.5MPa to 9.5MPa, the liquid hourly volume space velocity is 0.4h-1 -1~4.8h -1 .

[0024] In the above hydrocracking method, the outlet temperature of the hydropre-refining reaction zone is controlled to be above 380°C, and further to be between 381°C and 401°C.

[0025] In the above hydrocracking method, the main process conditions of the hydrocracking reaction zone are: average reaction temperature 300°C to 440°C, inlet hydrogen to oil volume ratio 550:1 to 1200:1, operating pressure 6.0MPa to 10.0MPa, liquid hourly volume space velocity 0.3h-1 to 5.0h-1; preferably, average reaction temperature 310°C to 435°C, inlet hydrogen to oil volume ratio 600:1 to 1150:1, operating pressure 6.5MPa to 9.5MPa, liquid hourly volume space velocity 0.4h-1 -1 ~4.8h -1 .

[0026] In the above-mentioned hydrocracking method, the number of defined catalyst beds is consistent with the prior art, that is, the catalyst for hydrocracking is loaded on different catalyst beds divided by a cold hydrogen box in the reaction zone, and different catalyst beds are defined in sequence according to the flow direction of the reactant flow. The cold hydrogen box has its own structure and function, but for the present invention, its segmentation effect is emphasized. In the diesel hydrocracking unit, according to the distribution of temperature rise and other conditions, generally only 3 catalyst beds of equal height are designed. When producing heavy naphtha, theoretically isothermal operation can maximize its output, but it cannot be achieved industrially. It can be adjusted to this trend by lowering the hot spot temperature, that is, increasing the number of catalyst beds. The more the hot spot temperature is reduced, the closer it is to the isothermal reaction operation of the cracking section, the lower the gasification rate will be, and the heavy stone selectivity will be improved. However, in view of the actual application situation, the present invention only points out its advantages and does not carry out unlimited bed design.

[0027] In the above-mentioned hydrocracking method, the operations such as product separation and fractionation mentioned are conventional processes in the art and belong to the category of physical flash evaporation, which will not be described again here.

[0028] In the above hydrocracking method, the main components of liquefied gas are carbon three and normal isomerized carbon four, and light naphtha is a liquid phase fraction in the fractionation product with a distillation temperature between liquefied gas and heavy naphtha (preferably its dry point is <65°C). After the introduction of the two, the probability of further cracking and lightening of heavy naphtha can be reduced, thereby achieving the inhibition of equilibrium reaction, and the loss of the two in the circulation process is extremely low. Here, the feeding of the circulating logistics in sections (diesel for refining; liquefied gas and light naphtha for cracking) is conducive to avoiding further accumulation of polycyclic aromatic hydrocarbons in the circulating diesel fraction. Based on the difference between the output mass of heavy naphtha and the mass of coker gasoline added, the circulating amount of liquefied gas accounts for 6wt% to 11wt% of the difference; the circulating amount of light naphtha accounts for 12wt% to 18wt% of the difference.

[0029] Compared with the prior art, it has the following beneficial effects:

[0030] (1) The hydrocracking method of the present invention achieves component dispersion by retaining high aromatic components during the full cycle reaction when processing diesel raw materials, especially straight-run diesel, which can reduce the tendency of adverse reactions to occur, and can improve the quality of heavy naphtha while increasing the production. In terms of process means, the characteristics of different types of raw materials and the characteristics of the reaction in the hydropre-refining process are integrated, and the temperature of the pre-refining section is increased by utilizing the reaction temperature rise through the reasonable introduction of high-quality target components, and the thermodynamic restrictions of aromatic saturation are combined to achieve the optimal retention of aromatics. The high-quality material selected for introduction is the coking gasoline component, which increases the production of heavy naphtha while making reasonable use of "low added value" resources. In addition, multiple hydrocracking catalyst beds are set in the hydrocracking zone to reduce the probability of side reactions occurring at high temperature points; and according to the positive carbon ion reaction mechanism of the hydrocracking reaction zone, the light components are recycled starting from the theoretically minimum component carbon three generated during the reaction process. The return of liquefied gas and light stone to the cracking reaction zone achieves uniform suppression of the entire fraction of the reaction equilibrium. The circulating diesel fraction is returned to the pre-refining section, which will reduce the accumulation of polycyclic aromatic hydrocarbons that are very likely to occur in the medium and low pressure reaction process without affecting the reaction process, which is beneficial to the long-term operation of the device. The method of the present invention can achieve the purpose of producing more high-quality heavy naphtha.

[0031] (2) In response to the product structure requirements of the current market, the present invention simultaneously processes two "low value-added" products, straight-run diesel and coker gasoline, and successfully improves their outlets. At the same time, the converted heavy naphtha is a high-quality chemical product, which is a typical "rational utilization of resources". Whether it is a new device or an existing device, it is easy to implement and the effect is also significant. It is an indispensable technology for enterprise transformation and upgrading. BRIEF DESCRIPTION OF THE DRAWINGS

[0032] Figure 1 It is a schematic diagram of the process flow of the present invention. DETAILED DESCRIPTION

[0033] The method of the present invention is described below in conjunction with the accompanying drawings. Some conventional unit diagrams that are not necessary for description are integrated and summarized as a separation system and a fractionation system.

[0034] like Figure 1 As shown, the process flow of the present invention is as follows:

[0035] The diesel feedstock 1, fractionated circulating diesel 15 and circulating hydrogen 8 are mixed and enter the upper part of the hydrogenation pre-refining reaction zone 2 together. The coking gasoline feedstock 3 enters the middle and lower part of the hydrogenation pre-refining reaction zone 2. After the hydrogenation pre-refining reaction, the effluent 4 is mixed with part of the circulating liquefied gas 12 and light naphtha and circulating hydrogen 8 and enters the hydrocracking reaction zone 5 for reaction. The effluent 6 enters the separation system 7, and the circulating hydrogen 8 is obtained at the top. The liquid phase separation effluent 9 at the bottom enters the fractionation system 10, and the dry gas 11, liquefied gas 12, light naphtha 13, heavy naphtha 14 and circulating diesel 15 are discharged from top to bottom in sequence; the liquefied gas 12 and the light naphtha 13 are at least partially recycled back to the hydrocracking reaction zone 5, and the diesel 15 is recycled back to the hydrogenation pre-refining reaction zone 2.

[0036] The method of the present invention is further described below with reference to the embodiments and comparative examples.

[0037] In the present invention, unless otherwise specified, percentages refer to mass fractions.

[0038] According to the method described in the present invention, some embodiments and comparative examples are carried out to examine the effect of producing more heavy naphtha. The processes adopted are all extreme single-stage full-cycle processes, the purpose of which is to produce more heavy naphtha. The reference for the proportion of liquefied gas and light naphtha is the difference between the total output of heavy naphtha and the amount of coker gasoline blended.

[0039] The process of the method of the present invention is as follows: Figure 1 The properties of the raw oil involved are shown in Table 1, the main physical and chemical properties of the catalyst are shown in Table 2, the relevant control conditions are shown in Table 3, and the effect comparison is shown in Table 4.

[0040] Example 1

[0041] This embodiment is based on the method and Figure 1 The process operation is to process the straight-run diesel feedstock in Table 1, the hydropre-refining reaction zone is filled with FF-66 in Table 2, and the hydrocracking reaction zone is filled with FC-50 in Table 2. The operating conditions of the hydropre-refining reaction zone are as follows: the liquid hourly volume space velocity of the straight-run diesel feedstock is 2.0h -1The reaction temperature is 344°C, the reaction pressure is 9.5 MPa, the inlet hydrogen-oil volume ratio of the straight-run diesel feedstock is 600:1, and the coking gasoline in Table 1 is injected at a height of 30% from the bottom of the catalyst bed in the pre-refining reaction zone, and the injection amount is 30% of the mass of the straight-run diesel feedstock. The outlet temperature of the hydrogenation pre-refining reaction zone is 390°C. The hydrogenation pre-refining product oil obtained after the raw oil reacts in the hydrogenation pre-refining reaction zone enters the top of the hydrocracking reactor for hydrocracking reaction. Four hydrocracking catalyst beds are set in the hydrocracking reaction zone. The reaction conditions are liquid hourly volume space velocity of 2.0h -1 , the reaction temperature is 367℃, the reaction pressure is 9.2MPa, the hydrogen-oil volume ratio is 800:1, the single-pass conversion rate is 60%, and the generated oil is subjected to subsequent separation and fractionation operations to obtain dry gas, liquefied gas, light naphtha, heavy naphtha and circulating diesel. The circulating diesel returns to the hydrogenation pre-refining reaction zone to continue the reaction, and part of the liquefied gas and light naphtha returns to the hydrocracking reaction zone to continue the reaction. The difference between the output quality of heavy naphtha and the quality of coker gasoline blending is taken as the benchmark, of which the circulating liquefied gas accounts for 8% of the difference and the circulating light naphtha accounts for 14% of the difference.

[0042] Example 2

[0043] This embodiment is based on the method and Figure 1 The process operation is to process the straight-run diesel feedstock in Table 1, the hydropre-refining reaction zone is filled with FF-66 in Table 2, and the hydrocracking reaction zone is filled with FC-50 in Table 2. The operating conditions of the hydropre-refining reaction zone are as follows: the liquid hourly volume space velocity of the straight-run diesel feedstock is 2.0h -1 The reaction temperature is 348°C, the reaction pressure is 7.5MPa, the inlet hydrogen-oil volume ratio of the straight-run diesel feedstock is 600:1, and the coking gasoline in Table 1 is injected at a height of 30% from the bottom of the catalyst bed in the pre-refining reaction zone, and the injection amount is 30% of the mass of the straight-run diesel feedstock. The outlet temperature of the hydrogenation pre-refining reaction zone is 385°C. The hydrogenation pre-refining product oil obtained after the raw oil reacts in the hydrogenation pre-refining reaction zone enters the top of the hydrocracking reactor for hydrocracking reaction. Four hydrocracking catalyst beds are set in the hydrocracking reaction zone. The reaction conditions are liquid hourly volume space velocity of 2.0h -1 , the reaction temperature is 371℃, the reaction pressure is 7.2MPa, the hydrogen-oil volume ratio is 800:1, the single-pass conversion rate is 60%, and the generated oil is subjected to subsequent separation and fractionation operations to obtain dry gas, liquefied gas, light naphtha, heavy naphtha and circulating diesel. The circulating diesel returns to the hydrogenation pre-refining reaction zone to continue the reaction, and part of the liquefied gas and light naphtha returns to the hydrocracking reaction zone to continue the reaction. The difference between the output quality of heavy naphtha and the quality of coker gasoline blending is taken as the benchmark, of which the circulating liquefied gas accounts for 8% of the difference and the circulating light naphtha accounts for 14% of the difference.

[0044] Example 3

[0045] This embodiment is based on the method and Figure 1 The process operation is to process the straight-run diesel feedstock in Table 1, the hydropre-refining reaction zone is filled with FF-66 in Table 2, and the hydrocracking reaction zone is filled with FC-50 in Table 2. The operating conditions of the hydropre-refining reaction zone are as follows: the liquid hourly volume space velocity of the straight-run diesel feedstock is 2.0h -1 , the reaction temperature is 344°C, the reaction pressure is 9.5MPa, the inlet hydrogen-oil volume ratio of the straight-run diesel feedstock is 600:1, the coking gasoline in Table 1 is injected at a height of 30% from the bottom of the catalyst bed in the pre-refining reaction zone, and the injection amount is 30% of the mass of the straight-run diesel feedstock, the outlet temperature of the hydrogenation pre-refining reaction zone is 395°C, the hydrogenation pre-refining product oil obtained after the raw oil reacts in the hydrogenation pre-refining reaction zone enters the top of the hydrocracking reactor for hydrocracking reaction, and 4 hydrocracking catalyst beds are set in the hydrocracking reaction zone. The reaction conditions are liquid hourly volume space velocity of 2.0h -1 , the reaction temperature is 366℃, the reaction pressure is 9.2MPa, the hydrogen-oil volume ratio is 800:1, the single-pass conversion rate is 60%, and the generated oil is subjected to subsequent separation and fractionation operations to obtain dry gas, liquefied gas, light naphtha, heavy naphtha and circulating diesel. The circulating diesel returns to the hydrogenation pre-refining reaction zone to continue the reaction, and part of the liquefied gas and light naphtha returns to the hydrocracking reaction zone to continue the reaction. The difference between the output quality of heavy naphtha and the quality of coker gasoline blending is taken as the benchmark, of which the circulating liquefied gas accounts for 8% of the difference and the circulating light naphtha accounts for 14% of the difference.

[0046] Example 4

[0047] This embodiment is based on the method and Figure 1 The process operation is to process the straight-run diesel feedstock in Table 1, the hydropre-refining reaction zone is filled with FF-66 in Table 2, and the hydrocracking reaction zone is filled with FC-50 in Table 2. The operating conditions of the hydropre-refining reaction zone are as follows: the liquid hourly volume space velocity of the straight-run diesel feedstock is 2.0h -1 The reaction temperature is 344°C, the reaction pressure is 9.5 MPa, the inlet hydrogen-oil volume ratio of the straight-run diesel feedstock is 600:1, and the coking gasoline in Table 1 is injected at a height of 20% from the bottom of the catalyst bed in the pre-refining reaction zone, and the injection amount is 25% of the mass of the straight-run diesel feedstock. The outlet temperature of the hydrogenation pre-refining reaction zone is 390°C. The hydrogenation pre-refining product oil obtained after the raw oil reacts in the hydrogenation pre-refining reaction zone enters the top of the hydrocracking reactor for hydrocracking reaction. Five hydrocracking catalyst beds are set in the hydrocracking reaction zone. The reaction conditions are liquid hourly volume space velocity of 1.8 h -1, the reaction temperature is 364℃, the reaction pressure is 9.2MPa, the hydrogen-oil volume ratio is 800:1, the single-pass conversion rate is 60%, and the generated oil is subjected to subsequent separation and fractionation operations to obtain dry gas, liquefied gas, light naphtha, heavy naphtha and circulating diesel. The circulating diesel returns to the hydrogenation pre-refining reaction zone to continue the reaction, and part of the liquefied gas and light naphtha returns to the hydrocracking reaction zone to continue the reaction. The difference between the output quality of heavy naphtha and the quality of coker gasoline blending is taken as the benchmark, of which the circulating liquefied gas accounts for 9% of the difference and the circulating light naphtha accounts for 15% of the difference.

[0048] Comparative Example 1

[0049] This comparative example is operated according to the conventional single-stage tail oil full circulation process to process the straight-run diesel and coker gasoline raw materials in Table 1, wherein the coker gasoline is 30% of the straight-run diesel, the hydropre-refining reaction zone is filled with FF-66 in Table 2, and the hydrocracking reaction zone is filled with FC-50 in Table 2. The straight-run diesel and coker gasoline raw materials and hydrogen are fed from the top of the hydropre-refining reaction zone. For the straight-run diesel raw material, the liquid hourly volume space velocity in the hydropre-refining reaction zone is 2.0 h -1 The reaction temperature is 344°C, the reaction pressure is 9.5 MPa, the inlet hydrogen-oil volume ratio is 600:1, and the hydropre-refined oil obtained after the feed oil reacts in the hydropre-refining reaction zone enters the top of the hydrocracking reactor for hydrocracking reaction. Four hydrocracking catalyst beds are set in the hydrocracking reaction zone. The reaction conditions are liquid hourly volume space velocity of 2.0 h -1 The reaction temperature is 367°C, the reaction pressure is 9.2MPa, the hydrogen-oil volume ratio is 800:1, the single-pass conversion rate is 60%, and the generated oil is subjected to subsequent separation and fractionation operations to obtain dry gas, liquefied gas, light naphtha, heavy naphtha and circulating diesel. The circulating diesel returns to the hydrogenation pre-refining reaction zone to continue the reaction.

[0050] Comparative Example 2

[0051] In this comparative example, the straight-run diesel feedstock in Table 1 is processed, the hydropre-refining reaction zone is filled with FF-66 in Table 2, and the hydrocracking reaction zone is filled with FC-50 in Table 2. The straight-run diesel feedstock and hydrogen are fed from the top of the hydropre-refining reaction zone. The operating conditions of the hydropre-refining reaction zone are as follows: the liquid hourly volume space velocity of the straight-run diesel feedstock is 2.0 h -1The reaction temperature is 348°C, the reaction pressure is 7.5MPa, the inlet hydrogen-oil volume ratio of the straight-run diesel feedstock is 600:1, and the coking gasoline in Table 1 is injected at a height of 30% from the bottom of the catalyst bed in the pre-refining reaction zone, and the injection amount is 30% of the mass of the straight-run diesel feedstock. The outlet temperature of the hydrogenation pre-refining reaction zone is 385°C. The hydrogenation pre-refining product oil obtained after the raw oil reacts in the hydrogenation pre-refining reaction zone enters the top of the hydrocracking reactor for hydrocracking reaction. Four hydrocracking catalyst beds are set in the hydrocracking reaction zone. The reaction conditions are liquid hourly volume space velocity of 2.0h -1 , reaction temperature is 371℃, reaction pressure is 7.2MPa, hydrogen-oil volume ratio is 800:1, single-pass conversion rate is 60%, and the generated oil is subjected to subsequent separation and fractionation operations to obtain dry gas, liquefied gas, light naphtha, heavy naphtha and circulating diesel. The circulating diesel and part of the liquefied gas and light naphtha are returned to the hydrocracking reaction zone to continue the reaction. The difference between the output quality of heavy naphtha and the quality of coker gasoline blending is taken as the benchmark, of which the circulating liquefied gas accounts for 8% of the difference and the circulating light naphtha accounts for 14% of the difference.

[0052] Comparative Example 3

[0053] In this comparative example, the straight-run diesel feedstock in Table 1 is processed, the hydropre-refining reaction zone is filled with FF-66 in Table 2, and the hydrocracking reaction zone is filled with FC-50 in Table 2. The straight-run diesel feedstock and hydrogen are fed from the top of the hydropre-refining reaction zone. The operating conditions of the hydropre-refining reaction zone are as follows: the liquid hourly volume space velocity of the straight-run diesel feedstock is 2.0 h -1 The reaction temperature is 348°C, the reaction pressure is 7.5MPa, the inlet hydrogen-oil volume ratio of the straight-run diesel feedstock is 600:1, and the coking gasoline in Table 1 is injected at a height of 30% from the bottom of the catalyst bed in the pre-refining reaction zone, and the injection amount is 30% of the mass of the straight-run diesel feedstock. The outlet temperature of the hydrogenation pre-refining reaction zone is 385°C. The hydrogenation pre-refining product oil obtained after the raw oil reacts in the hydrogenation pre-refining reaction zone enters the top of the hydrocracking reactor for hydrocracking reaction. Four hydrocracking catalyst beds are set in the hydrocracking reaction zone. The reaction conditions are liquid hourly volume space velocity of 2.0h -1 , the reaction temperature is 371℃, the reaction pressure is 7.2MPa, the hydrogen-oil volume ratio is 800:1, the single-pass conversion rate is 60%, and the generated oil is subjected to subsequent separation and fractionation operations to obtain dry gas, liquefied gas, light naphtha, heavy naphtha and circulating diesel. The circulating diesel returns to the hydrocracking reaction zone to continue the reaction, and part of the liquefied gas and light naphtha returns to the hydrogenation pre-refining reaction zone to continue the reaction. The difference between the output quality of heavy naphtha and the quality of coker gasoline blending is taken as the benchmark, of which the circulating liquefied gas accounts for 8% of the difference and the circulating light naphtha accounts for 14% of the difference.

[0054] Comparative Example 4

[0055] This comparative example is based on the method of the present invention and Figure 1 The process operation is to process the straight-run diesel feedstock in Table 1, the hydropre-refining reaction zone is filled with FF-66 in Table 2, and the hydrocracking reaction zone is filled with FC-50 in Table 2. The operating conditions of the hydropre-refining reaction zone are as follows: the liquid hourly volume space velocity of the straight-run diesel feedstock is 2.0h -1 , the reaction temperature is 348°C, the reaction pressure is 7.5MPa, the inlet hydrogen-oil volume ratio of the straight-run diesel feedstock is 600:1, and the coking gasoline in Table 1 is injected at a height of 30% of the total height from the bottom of the catalyst bed in the pre-refining reaction zone, and the injection amount is 30% of the mass of the straight-run diesel feedstock. The outlet temperature of the hydrogenation pre-refining reaction zone is 385°C. The hydrogenation pre-refining product oil obtained after the raw oil reacts in the hydrogenation pre-refining reaction zone enters the top of the hydrocracking reactor for hydrocracking reaction. Four hydrocracking catalyst beds are set in the hydrocracking reaction zone. The reaction conditions are liquid hourly volume space velocity of 2.0h -1 , the reaction temperature is 371℃, the reaction pressure is 7.2MPa, the hydrogen-oil volume ratio is 800:1, the single-pass conversion rate is 60%, and the generated oil is subjected to subsequent separation and fractionation operations to obtain dry gas, liquefied gas, light naphtha, heavy naphtha and circulating diesel. The circulating diesel returns to the hydrocracking reaction zone to continue the reaction, and part of the liquefied gas and light naphtha returns to the hydrogenation pre-refining reaction zone to continue the reaction. The difference between the output quality of heavy naphtha and the quality of coker gasoline blending is taken as the benchmark, of which the circulating liquefied gas accounts for 4% of the difference and the circulating light naphtha accounts for 7% of the difference.

[0056] Table 1 Raw oil properties

[0057] Source of raw materials Straight run diesel Coking gasoline <![CDATA[Density (20 °C), g·cm -3 > 0.8769 0.7211 Distillation range, ℃ 209～368 37～177 <![CDATA[Sulfur, mg·kg -1 > 8100 7000 <![CDATA[Nitrogen, mg·kg -1 > 150 196 Olefin content, wt% -- 41.9

[0058] Table 2 Main physical and chemical properties of catalysts

[0059] project FF-66 FC-50 Chemical composition Mo-Ni W-Ni Physical properties Appearance and shape Shamrock Strips Cylindrical Bar Crushing strength, N / cm ≥150 ≥150 Particle diameter, mm 1.0～1.4 1.4～1.8 Molecular sieve content, wt% -- Y-type molecular sieve / 20

[0060] Table 3 Main reaction conditions for each example

[0061]

[0062] Table 4 Comparison of the effects obtained in each case

[0063]

Claims

1. A hydrocracking method for preparing heavy naphtha, the method comprising a hydropre-refining reaction zone and a hydrocracking reaction zone, the process comprising: The diesel feedstock, hydrogen and circulating oil are mixed and fed from the top of the hydrogenation pre-refining reaction zone, and the coker gasoline feedstock is fed from the middle and lower part of the hydrogenation pre-refining reaction zone. After the hydrogenation refining reaction, the effluent of the hydrogenation pre-refining reaction zone is mixed with at least part of the liquefied gas and at least part of the light naphtha and then enters the hydrocracking reaction zone for reaction. The effluent of the hydrocracking reaction zone is separated and fractionated to obtain dry gas, liquefied gas, light naphtha, heavy naphtha and diesel, wherein the diesel is circulated back to the hydrogenation pre-refining reaction zone as the circulating oil, and at least part of the liquefied gas and at least part of the light naphtha are circulated back to the hydrocracking reaction zone; The feed mass ratio of the diesel raw material to the coker gasoline raw material is 8:1 to 2:1; Based on the difference between the output mass of heavy naphtha and the added mass of coker gasoline, the circulating volume of liquefied gas accounts for 6wt%~11wt% of the difference, and the circulating volume of light naphtha accounts for 12wt%~18wt% of the difference; The hydrocracking reaction zone is provided with more than three hydrocracking catalyst beds.

2. The method according to claim 1, characterized in that The middle and lower part of the hydrogenation pre-refining reaction zone refers to: the height of the middle and lower feeding position from the bottom of the last catalyst bed in the hydrogenation pre-refining reaction zone accounts for 15% to 50% of the total height of the reaction zone, wherein the total height of the reaction zone is measured by the total height of the catalyst loaded in the hydrogenation pre-refining reaction zone.

3. The method according to claim 2, characterized in that The middle and lower part of the hydrogenation pre-refining reaction zone refers to: the height of the middle and lower feeding position from the bottom of the last catalyst bed in the hydrogenation pre-refining reaction zone accounts for 20% to 45% of the total height of the reaction zone.

4. The method according to claim 1, characterized in that: The feed mass ratio of the diesel feedstock to the coking gasoline feedstock is 6:1 to 3:

1.

5. The method according to claim 1, characterized in that The hydrogenation pre-refining reaction zone is filled with a refined catalyst, and based on the weight of the catalyst, the content of the VIB group metal component in terms of metal oxide is 9wt%~35wt%; the content of the VIII group metal in terms of metal oxide is 0.5wt%~7wt%.

6. The method according to claim 5, characterized in that The hydrogenation pre-refining reaction zone is filled with a refined catalyst, and based on the weight of the catalyst, the content of the VIB group metal component in terms of metal oxide is 10wt%~30wt%; the content of the VIII group metal in terms of metal oxide is 1.1wt%~6wt%.

7. The method according to claim 1, characterized in that The hydrocracking reaction zone is filled with a hydrocracking catalyst, wherein the content of the hydrogenation metal in terms of oxide is 12wt% to 33wt% based on the weight of the catalyst, and the content of the molecular sieve is 1wt% to 30wt%; the hydrogenation metal is selected from one or more of Group VIB, Group VIIB, Group VIII metals, metal oxides or metal sulfides.

8. The method according to claim 7, characterized in that The hydrocracking reaction zone is filled with a hydrocracking catalyst, and the content of the hydrogenation metal in terms of oxide is 14 wt% to 32 wt% based on the weight of the catalyst.

9. The method according to claim 1, characterized in that: The diesel raw material is one or more of straight-run diesel, ebullated-bed hydrogenated diesel, catalytic diesel and coking diesel.

10. The method according to claim 9, characterized in that The diesel feedstock is a diesel feedstock containing straight-run diesel, and the straight-run diesel accounts for more than 50wt% of the total mass of the diesel feedstock.

11. The method according to claim 10, characterized in that The straight-run diesel accounts for 51 wt% to 100 wt% of the total mass of the diesel raw material.

12. The method according to claim 9, characterized in that The straight-run diesel has a nitrogen content of less than 0.06 wt %, a distillation point of 155-255° C., a final distillation point of 315-385° C., and an aromatics content of 5 wt %-35 wt %.

13. The method according to claim 12, characterized in that The nitrogen content in the straight-run diesel is below 0.03wt%, and the aromatics content is 6wt%-34wt%.

14. The method according to claim 13, characterized in that The nitrogen content in the straight-run diesel is 0.001wt%~0.029wt.

15. The method according to claim 1, characterized in that The coker gasoline has an initial distillation point of 22-45° C., a final distillation point of 160-205° C., and an olefin content of more than 30 wt%.

16. The method according to claim 15, characterized in that The coking gasoline has an initial boiling point of 25-40° C., a final boiling point of 170-200° C., and an olefin content of more than 35wt%.

17. The method according to claim 15, characterized in that The olefin content of the coker gasoline is 36 wt % to 70 wt %.

18. The method according to claim 1, characterized in that The effluent from the hydrocracking reaction zone is separated by a separation system, and hydrogen is separated from the top of the separation system and circulated back to the hydrogenation pre-refining reaction zone and the hydrocracking reaction zone as circulating hydrogen.

19. The method according to claim 1, characterized in that For the diesel feedstock at the inlet of the hydrogenation pre-refining reaction zone, the process conditions of the hydrogenation pre-refining reaction zone are as follows: average reaction temperature 260℃~390℃, inlet hydrogen-oil volume ratio 350:1~900:1, operating pressure 6.0MPa~10.0MPa, liquid hourly volume space velocity 0.3h-1~5.0h -1 .

20. The method according to claim 19, characterized in that For the diesel feedstock at the inlet of the hydrogenation pre-refining reaction zone, the process conditions of the hydrogenation pre-refining reaction zone are as follows: average reaction temperature 270℃~370℃, inlet hydrogen-oil volume ratio 400:1~850:1, operating pressure 6.5MPa~9.5MPa, liquid hourly volume space velocity 0.4h -1 ~4.8h -1 .

21. The method according to claim 1, characterized in that The outlet temperature of the hydrogenation pre-refining reaction zone is controlled to be above 380°C.

22. The method according to claim 21, characterized in that The outlet temperature of the hydrogenation pre-refining reaction zone is controlled to be 381°C~401°C.

23. The method according to claim 1, characterized in that The hydrocracking reaction zone is provided with 4 to 6 hydrocracking catalyst beds.

24. The method according to claim 1, characterized in that The process conditions of the hydrocracking reaction zone are: average reaction temperature 300℃~440℃, inlet hydrogen-oil volume ratio 550:1~1200:1, operating pressure 6.0MPa~10.0MPa, liquid hourly volume space velocity 0.3h-1~5.0h -1 .

25. The method according to claim 24, characterized in that The process conditions of the hydrocracking reaction zone are: average reaction temperature 310℃~435℃, inlet hydrogen-oil volume ratio 600:1~1150:1, operating pressure 6.5MPa~9.5MPa, liquid hourly volume space velocity 0.4h -1 ~4.8h -1 .

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