Adjusting method for switching wax oil hydrocracking to catalytic cracking diesel hydrocracking

Abstract

The invention discloses an adjusting method for switching wax oil hydrocracking into catalytic cracking diesel hydrocracking, which comprises the following steps of: (1) reducing the average temperature of each bed layer of a wax oil hydrocracking device in stable operation to 90-230 DEG C, preferably 180-230 DEG C, and switching the fed material to direct current diesel; (2) when the sulfur content of the reaction effluent of the hydrocracking device is less than 500 [mu] g.g <-1 >, adjusting the reaction conditions of the hydrocracking device to enable the final boiling point of the reaction effluent to be 1t; keeping the temperature constant after the light oil mass yield of the fraction at 280 DEG C is 1-10%; (3) performing reduction treatment on the hydrocracking catalyst of each bed layer in the constant temperature process in the step (2), preferably, gradually increasing the reduction depth of the hydrocracking catalyst along the material flow direction; and (4) after the reduction treatment is finished, directly introducing catalytic cracking diesel oil or introducing the catalytic cracking diesel oil for reaction after vulcanization to obtain naphtha. The method can be switched to catalytic cracking diesel oil hydrocracking on line on a wax oil hydrocracking device, and naphtha with high aromatic potential and high yield is obtained.

Description

Technical Field

[0001] The present invention relates to an adjustment method for switching from wax oil hydrocracking to catalytic diesel hydrocracking, and in particular to an adjustment method for a smoothly operating wax oil hydrocracking unit when the raw materials to be processed are changed and catalytic diesel needs to be processed. Background Art

[0002] The hydrocracking process is a process in which heavy distillate oil (VGO, CGO, DAO) is hydrodesulfurized, hydrodenitrogenated, and polycyclic aromatic hydrocarbons are hydrosaturated and ring-opened cracked under the conditions of hydrogen, high temperature, high pressure and catalyst to convert it into target products such as light oil and middle distillate oil.

[0003] CN109988629 discloses a wax oil hydrocracking method and system, which includes: subjecting wax oil feedstock to hydrogenation pretreatment to obtain a hydropretreatment stream; contacting the hydropretreatment stream and a hydrogen-containing stream with a first hydrocracking catalyst to obtain a first hydrocracking stream, and dividing the first hydrocracking stream into two parts, first hydrocracking stream a and first hydrocracking stream b; contacting the first hydrocracking stream b and the hydrogen-containing stream with a second hydrocracking catalyst, and separating and fractionating the obtained second hydrocracking stream to obtain a hydrocracking tail oil product; contacting the first hydrocracking stream a, at least a portion of the hydrocracking tail oil product, the hydrogen-containing stream with a hydroisomerization cracking catalyst, and separating and fractionating the obtained hydroisomerization cracking stream. The method provided by this invention can meet the needs of producing naphtha products, aviation kerosene products, diesel products, and tail oil products of different specifications. CN110938466 discloses a wax oil hydrocracking method, comprising the following contents: (1) the wax oil raw material is mixed with hydrogen and first enters a hydrorefining reactor for desulfurization, denitrogenation and aromatic saturation reaction; (2) the effluent of step (1) enters a hydrocracking reactor, and the hydrocracking reactor comprises 1st to nth catalyst beds along the material direction, wherein n≥3, preferably n is 3 or 4, wherein the 1st to n-1th catalyst beds are filled with a hydrocracking catalyst containing a modified Y molecular sieve; and the nth catalyst bed is filled with a hydroisomerization catalyst containing a β and / or ZSM series molecular sieve with strong isomerization performance; (3) the effluent of the hydrocracking reactor is subjected to gas-liquid separation and liquid phase fractionation to obtain naphtha, jet fuel, diesel and tail oil. This method uses wax oil as raw material, and by setting a reasonable catalyst grading method and adjusting the catalyst composition and structure in the graded bed, the production of heavy naphtha with high aromatic potential is increased, and the quality of jet fuel and diesel products is improved. CN103102913 discloses a wax oil hydrotreating method for producing by-product diesel. First, bio-oil is mixed with recycled hydrogen for hydrogenation. The resulting liquid is then mixed with wax oil feedstock for hydrogenation and fractionated to produce naphtha, high-quality diesel, and hydrogenated wax oil products. This method can produce high-quality diesel as a by-product, has good catalyst activity and stability, and allows the device to operate stably over a long period of time. CN101875856 discloses a combined wax oil hydrotreating and catalytic cracking method. The hydrotreated feedstock undergoes a hydrogenation reaction in the presence of hydrogen and a hydrotreating catalyst. The liquid phase of the hydrogenation reaction effluent is directly fed into a catalytic cracking unit. The catalytic cracking heavy fraction, which is separated after the catalytic cracking gasoline, is mixed with the hydrotreated feedstock for hydrogenation. This method can maximize gasoline production while saving equipment investment, fully utilizing reaction heat, and reducing production energy consumption. CN115785994 discloses a wax oil hydrogenation method and apparatus. The method comprises: the wax oil feedstock is first hydrotreated to separate into gas and liquid. The liquid enters a hydrocracking reactor and is separated into a lighter fraction and a heavier fraction after passing through a liquid distribution assembly.The lighter part is not subjected to hydrocracking, while the heavier part is subjected to hydrocracking reaction, and the resulting products are separated to obtain naphtha fraction and diesel fraction. The method of the invention controls the reaction depth in a more balanced manner, ultimately achieving the purpose of increasing the yield of the target product, especially improving the selectivity and yield of naphtha, while eliminating the risk of "temperature runaway" or flooding of the bed. CN113430002 discloses a method for hydrogenating inferior wax oil. The method involves a direct graded combination of two different types of catalysts, namely, a type II catalyst containing an organic additive and a type I / II mixed catalyst. The inferior wax oil feed first flows through the type II catalyst containing an organic additive, and then the effluent contacts the type I / II mixed catalyst to achieve the purpose of removing sulfur and nitrogen impurities by coking depth, while improving the stability of the catalyst. CN 111100695 discloses a segmented feed wax oil hydroprocessing method, comprising the following steps: (1) under hydrorefining process conditions, vacuum wax oil is mixed with hydrogen and enters a first hydrorefining reaction zone, wherein a hydrorefining catalyst A is loaded; (2) the hydrorefining reaction effluent obtained in step (1) is mixed with secondary processed wax oil and enters a second hydrorefining reaction zone, wherein a hydrogenation protective agent, a hydrodemetallization catalyst, and a hydrorefining catalyst B are sequentially loaded in the second hydrorefining reaction zone along the flow direction; (3) the hydrorefining reaction effluent obtained in step (2) enters a separation system, where gas, naphtha, diesel, and tail oil are separated and the tail oil is used as feed for a catalytic cracking unit. This method effectively reduces the energy consumption of the unit by segmented processing of different inferior raw materials, while also extending the operating cycle of the unit. CN102899081 discloses a wax oil hydrotreating method. Feedstock oil is mixed with hydrogen and sequentially fed into three hydrogenation reaction zones. These zones are respectively loaded with a hydrogenation protective agent, a nickel-molybdenum-tungsten catalyst, and a cobalt-molybdenum catalyst. This method improves cracking performance and increases light oil yield.

[0004] CN1955261 discloses a method for hydrocracking low-quality catalytically cracked diesel. This method mixes low-quality catalytically cracked diesel with heavy hydrocracking feedstock, first hydrocracking it, and then subjecting the resulting intermediate distillate to two-stage hydrocracking to produce high-aromatic potential naphtha and a tail oil with a low BMCI value. CN1955262 discloses a method for hydrotreating low-quality catalytically cracked diesel. This method utilizes a two-stage hydrocracking process to treat low-quality catalytically cracked diesel. This method can produce heavy naphtha with high aromatic potential. CN103805245 discloses a hydrogenation method that combines hydrocracking and hydrodearomatization. The catalytic diesel is hydrorefined with hydrogen in a countercurrent flow; the refined oil is dearomatized with a precious metal catalyst; the wax oil is subjected to a hydrocracking pretreatment reaction with hydrogen; the pretreatment effluent is mixed with the dearomatization effluent and subjected to a hydrocracking reaction, and the resulting tail oil enters a dearomatization reactor. This method enhances the cracking function of the hydrocracking catalyst. CN106047404 discloses a combined process method for increasing the production of high-octane gasoline using inferior catalytic diesel. The method first subjects the catalytic diesel to a hydrogenation refining reaction, and then reacts the refined liquid product with a precious metal catalyst to ultimately achieve the purpose of producing more high-octane gasoline.

[0005] During the normal operation of the hydrocracking unit, in order to meet the needs of production and market, it is necessary to temporarily change the raw oil to produce products that meet the current market needs. Since the hydrocracking operating conditions of different raw materials vary greatly, it is generally necessary to stop the operation for processing, and the operation efficiency is relatively low. Summary of the Invention

[0006] In view of the shortcomings of the existing technology, the present invention provides an adjustment method for switching from wax oil hydrocracking to catalytic cracking diesel hydrocracking. The method can switch to catalytic cracking diesel hydrocracking online on a wax oil hydrocracking unit to obtain naphtha with high aromatic potential and high yield.

[0007] A method for adjusting the process of switching from wax oil hydrocracking to catalytically cracked diesel hydrocracking, comprising the following steps:

[0008] (1) The average temperature of each bed of the stable-operating wax oil hydrocracking unit is cooled to 90°C to 230°C, preferably 180°C to 230°C, and the feed is switched to direct-flow diesel;

[0009] (2) When the sulfur content of the hydrocracking unit reaction effluent is less than 500 μg·g -1 When the concentration is less than 100 μg·g -1 , most preferably less than 10 μg·g -1 When the reaction conditions of the hydrocracking unit are adjusted, the mass yield of the light oil of the fraction with the final boiling point of the reaction effluent <280°C is 1% to 10%, and then the temperature is kept constant;

[0010] (3) During the constant temperature process in step (2), the hydrocracking catalyst in each bed is subjected to reduction treatment, preferably with the reduction depth of the hydrocracking catalyst gradually increasing along the direction of the flow;

[0011] (4) After the reduction treatment is completed, catalytic cracking diesel is directly introduced or catalytic cracking diesel is introduced after sulfurization to react to obtain naphtha with high yield and high aromatic potential.

[0012] In step (1) of the method of the present invention, the wax oil hydrocracking unit is provided with N hydrocracking catalyst beds, wherein N is at least 2, preferably 2 to 6, and most preferably 3 to 4; the packing volume ratio of the Nth hydrocracking catalyst bed to the N-1th catalyst bed is 1:0.1 to 1:10, preferably 1:0.5 to 1:5.0.

[0013] In step (1) of the method of the present invention, the wax oil feedstock contacts and reacts with the hydrocracking catalyst during stable operation, and the effluent of the wax oil hydrocracking reaction is separated and fractionated to obtain wax oil hydrocracking products including light naphtha, heavy naphtha, aviation kerosene, diesel and tail oil.

[0014] In step (1) of the method of the present invention, the hydrocracking catalyst loaded into the wax oil hydrocracking unit can be a commercially available product or prepared by conventional methods, or a regenerated hydrocracking catalyst can be used.

[0015] In step (1) of the method of the present invention, the cracking component of the hydrocracking catalyst loaded into the wax oil hydrocracking unit includes, but is not limited to, Y-type or β molecular sieves. The hydrocracking catalyst comprises a hydrogenation-active metal, a molecular sieve component, and an alumina carrier. The hydrogenation-active metal includes, but is not limited to, one or a mixture of Wo, Mo, Co, and Ni. The hydrocracking catalyst comprises a carrier and a supported hydrogenation metal. Based on the weight of the catalyst, the hydrogenation metal generally comprises a Group VIB metal component of the periodic table, such as tungsten and / or molybdenum, calculated as oxide, in an amount of 10% to 45%, preferably 15% to 30%; a Group VIII metal such as nickel and / or cobalt, calculated as oxide, in an amount of 1% to 7%, preferably 1.5% to 6.5%. The carrier is a molecular sieve and alumina, with emphasis on the molecular sieve content, generally 2% to 50%.

[0016] In step (1) of the method of the present invention, the wax oil is a vacuum wax oil, and the final boiling point of the vacuum wax oil is generally 450-550°C, preferably 460-520°C, the density is generally below 0.92g / cm3, preferably below 0.91g / cm3, the sulfur content is generally above 0.1wt%, preferably 0.1-3.0wt%, there is no special requirement for the nitrogen content, the aromatics content is generally below 60wt%, preferably 1.0-50wt%, and can be selected from various vacuum wax oils (VGO) obtained by processing Middle Eastern crude oil, such as one or more of Iranian VGO, Saudi VGO, etc.

[0017] In step (1) of the method of the present invention, the reaction temperature of the wax oil hydrocracking is 250°C to 450°C, preferably 310°C to 390°C; the operating pressure is 2.0MPa to 20.0MPa, preferably 8MPa to 16.0MPa; the hydrogen-to-oil volume ratio is 100:1 to 2500:1, preferably 500:1 to 1500:1; the volume space velocity is 0.1 to 5.0h -1 , preferably 0.5~3.0h -1 .

[0018] In step (1) of the method of the present invention, the initial boiling point of the straight-run diesel is generally greater than 170°C, preferably 170-280°C; the final boiling point is generally 210-410°C, preferably 260-390°C; the density is generally 0.65-0.90 g / cm 3 , preferably 0.70~0.85g / cm 3 ; Sulfur content is generally 500μg·g -1 Below, preferably 100 μg·g -1 The straight-run diesel can be selected from one or more of the following: low-sulfur aviation fuel or low-sulfur diesel obtained through hydrocracking reaction and low-sulfur straight-run diesel or low-sulfur straight-run aviation fuel obtained through processing.

[0019] In step (2) of the method of the present invention, the reaction conditions of the hydrocracking unit are adjusted so that the mass yield of light oil in the fraction with a final boiling point of the reaction effluent <280°C is 1% to 10%, which generally means adjusting the average reaction temperature of each bed to 230°C to 450°C, preferably 260°C to 390°C.

[0020] In step (3) of the method of the present invention, the reduction treatment adopts the method of introducing hydrogen into each bed layer, and the reduction degree of the catalyst in different beds can be controlled by adjusting and controlling the hydrogen injection time, hydrogen-to-oil ratio, hydrogen purity and H2S gas content in hydrogen. For example, hydrogen can be injected into each bed layer at the same time, and the hydrogen-to-oil volume ratio, hydrogen purity and H2S gas content in hydrogen of each bed layer can be changed without changing the hydrogen injection time; or the hydrogen injection time, hydrogen purity and H2S gas content in hydrogen of each bed layer can be changed without changing the hydrogen-to-oil volume ratio of each bed layer; or the hydrogen injection time, hydrogen-to-oil volume ratio and H2S gas content in hydrogen of each bed layer can be changed without changing the hydrogen purity of each bed layer; or the hydrogen injection time, hydrogen-to-oil volume ratio and hydrogen purity of each bed layer can be changed without changing the H2S gas content in hydrogen; or the hydrogen injection time, hydrogen-to-oil volume ratio, hydrogen purity and H2S gas content in hydrogen of each bed layer can be changed simultaneously to adjust the reduction degree of the catalyst in each bed layer.

[0021] In step (3) of the method of the present invention, the hydrogen purity of the injected hydrogen is 30% to 100%, preferably 65% ​​to 100%, and the amount of H2S gas contained in the injected hydrogen is 0.1 to 1000 μl·L-1, preferably 1 to 100 μl·L-1.

[0022] In step (3) of the method of the present invention, the adjustment of the hydrogen injection time of each bed layer is to adjust the hydrogen injection time of the hydrogen injection equipment of each bed layer respectively. The hydrogen injection time of a single bed layer is 0.1h to 50h, preferably 1h to 20h.

[0023] In step (3) of the method of the present invention, the hydrogen-to-oil volume ratio of each bed is adjusted by adjusting the hydrogen injection amount of the hydrogen injection equipment of each bed to achieve the purpose of flexibly adjusting the hydrogen-to-oil ratio of each bed. The hydrogen-to-oil ratio of a single bed is 100:1 to 5000:1, preferably 300:1 to 2500:1.

[0024] A non-limiting reduction treatment method is to introduce fresh hydrogen into each bed layer along the logistics direction, with the amount of increase in adjacent beds being 10%-30%, and the reduction time being 0.1h to 50h.

[0025] In step (4) of the method of the present invention, the catalytic cracking diesel oil has a final boiling point of generally 210 to 450° C., preferably 260 to 390° C., a density of generally 0.85 to 0.99 g / cm 3 , preferably 0.93 to 0.97 g / cm 3 , a sulfur content of generally above 0.1 wt %, preferably 0.1 to 3.0 wt %, no special requirements for nitrogen content, an aromatics content of generally above 60 wt %, preferably 80 to 90 wt %, and can be selected from catalytic cracking diesel oil processed in a catalytic cracking unit.

[0026] In step (4) of the method of the present invention, the catalytic cracking diesel hydrocracking reaction conditions are as follows: reaction temperature is 250°C to 450°C, preferably 330°C to 420°C; operating pressure is 2.0MPa to 20.0MPa, preferably 5.0MPa to 10.0MPa; hydrogen-to-oil volume ratio is 100:1 to 2500:1, preferably 500:1 to 1500:1; volume space velocity is 0.1 to 5.0h-1, preferably 0.5 to 2.0h-1.

[0027] Compared with the prior art, the adjustment method of the present invention for switching from wax oil hydrocracking to catalytic cracking diesel hydrocracking can be switched online to catalytic cracking diesel hydrocracking on a normally operating wax oil hydrocracking unit without stopping the unit, thereby improving the effect of catalytic cracking diesel hydrocracking, especially the yield of heavy naphtha. DETAILED DESCRIPTION

[0028] The following examples and comparative examples are used to further illustrate the effects and effects of the method of the present invention. However, the following examples do not limit the method of the present invention. The percentages mentioned in the upper limit of the present invention are all mass percentages unless otherwise specified.

[0029] The hydrocracking catalyst used in these Examples and Comparative Examples was composed of a Y molecular sieve as the acidic component, W and Ni as the active metal components of the hydrocracking catalyst. Based on the weight of the hydrocracking catalyst, the Y molecular sieve content was 30%, the W content (calculated as oxide) was 25%, the Ni content (calculated as oxide) was 8%, the amorphous silica-alumina content was 26.18% by weight (a molar ratio of silicon oxide to aluminum oxide was 1.0:2.4), and the balance was small-pore alumina. The hydrocracking catalyst bed consisted of four layers. The properties of the wax oil, catalytically cracked diesel, and direct current diesel used in these Examples and Comparative Examples are shown in Table 1.

[0030] Table 1

[0031]

[0032] Comparative Example 1

[0033] In a stable operation of the wax oil hydrocracking unit, the wax oil hydrocracking reaction conditions are as follows: reaction temperature 375°C; operating pressure 14.7 MPa; hydrogen to oil volume ratio 1000:1; volume space velocity 1.0 h -1 Without introducing straight-run diesel or reduction treatment, catalytic cracking diesel was introduced directly. The operating conditions were as follows: reaction temperature 390°C; operating pressure 8 MPa; hydrogen-to-oil volume ratio 1200:1; volume space velocity 1.5 h -1 The heavy naphtha yield was 45% and the aromatic potential was 54.2.

[0034] Example 1

[0035] (1) The temperature of each bed of the stable operation wax oil hydrocracking unit was lowered to 180°C, and the feed was switched to direct current diesel. The wax oil hydrocracking reaction conditions were as follows: reaction temperature 375°C; operating pressure 14.7 MPa; hydrogen to oil volume ratio 1000:1; volume space velocity 1.0 h -1 .

[0036] (2) When the sulfur content of the hydrocracking unit reaction effluent is less than 500 μg·g -1 When the reaction conditions of the hydrocracking unit are adjusted so that the mass yield of the light oil of the fraction with the final boiling point of the reaction effluent < 210°C reaches 2%, the temperature is then kept constant;

[0037] (3) During the constant temperature process of step (2), the hydrocracking catalyst in each bed is subjected to hydrogen reduction treatment. Fresh hydrogen is introduced into each bed along the logistics direction. The hydrogen-to-oil ratio of fresh hydrogen introduced into the first catalyst is 300:1. The increase in the amount of hydrogen introduced into the adjacent beds is 10%. The reduction time is 1 hour.

[0038] (4) After the reduction treatment, catalytic cracking diesel was directly introduced. The operating conditions were as follows: reaction temperature 390°C; operating pressure 8 MPa; hydrogen-to-oil volume ratio 1200:1; volume space velocity 1.5 h -1 The heavy naphtha yield was 47.3% and the aromatic potential was 56.8%.

[0039] Example 2

[0040] (1) The temperature of each bed of the stable operation wax oil hydrocracking unit was lowered to 190°C, and the feed was switched to direct current diesel. The wax oil hydrocracking reaction conditions were as follows: reaction temperature 365°C; operating pressure 15.7 MPa; hydrogen to oil volume ratio 1200:1; volume space velocity 1.5 h -1 .

[0041] (2) When the sulfur content of the hydrocracking unit reaction effluent is less than 400 μg·g -1 When the reaction conditions of the hydrocracking unit are adjusted so that the mass yield of the light oil of the fraction with the final boiling point of the reaction effluent < 220°C reaches 3%, the temperature is then kept constant;

[0042] (3) During the constant temperature process of step (2), the hydrocracking catalyst in each bed is subjected to hydrogen reduction treatment. Fresh hydrogen is introduced into each bed along the logistics direction. The hydrogen-to-oil ratio of fresh hydrogen introduced into the first catalyst is 400:1. The increase in the amount of hydrogen introduced into the adjacent beds is 15%, and the reduction time is 4 hours.

[0043] (4) After the reduction treatment, catalytic cracking diesel was directly introduced. The operating conditions were as follows: reaction temperature 375°C; operating pressure 7.0 MPa; hydrogen-to-oil volume ratio 1000:1; volume space velocity 1.0 h -1 The heavy naphtha yield was 48.2% and the aromatic potential was 58.1%.

[0044] Example 3

[0045] (1) The temperature of each bed of the stable operation wax oil hydrocracking unit was lowered to 200°C, and the feed was switched to direct current diesel. The wax oil hydrocracking reaction conditions were as follows: reaction temperature 385°C; operating pressure 16.7 MPa; hydrogen to oil volume ratio 1300:1; volume space velocity 2.0 h -1 .

[0046] (2) When the sulfur content of the hydrocracking unit reaction effluent is less than 200 μg·g -1When the reaction conditions of the hydrocracking unit are adjusted so that the mass yield of the light oil of the fraction with the final boiling point of the reaction effluent <250°C reaches 1%, the temperature is then kept constant;

[0047] (3) During the constant temperature process of step (2), the hydrocracking catalyst in each bed is subjected to hydrogen reduction treatment. Fresh hydrogen is introduced into each bed along the logistics direction. The hydrogen-to-oil ratio of fresh hydrogen introduced into the first catalyst is 600:1. The increase in the amount of hydrogen introduced into the adjacent beds is 20%. The reduction time is 10 hours.

[0048] (4) After the reduction treatment, catalytic cracking diesel was directly introduced. The operating conditions were as follows: reaction temperature 395°C; operating pressure 6.5 MPa; hydrogen-to-oil volume ratio 1400:1; volume space velocity 1.5 h -1 The heavy naphtha yield was 50.7% and the aromatic potential was 61.3%.

[0049] Example 4

[0050] (1) The temperature of each bed of the stable operation wax oil hydrocracking unit was lowered to 220°C, and the feed was switched to direct current diesel. The wax oil hydrocracking reaction conditions were as follows: reaction temperature 405°C; operating pressure 13.2 MPa; hydrogen to oil volume ratio 1500:1; volume space velocity 0.5 h -1 .

[0051] (2) When the sulfur content of the hydrocracking unit reaction effluent is less than 50 μg·g -1 When the reaction conditions of the hydrocracking unit are adjusted so that the mass yield of the light oil of the fraction with the final boiling point of the reaction effluent <280°C reaches 3%, the temperature is then kept constant;

[0052] (3) During the constant temperature process of step (2), the hydrocracking catalyst in each bed is subjected to hydrogen reduction treatment. Fresh hydrogen is introduced into each bed along the logistics direction. The hydrogen-to-oil ratio of fresh hydrogen introduced into the first catalyst is 800:1. The increase in the amount of hydrogen introduced into the adjacent beds is 25%, and the reduction time is 15 hours.

[0053] (4) After the reduction treatment, catalytic cracking diesel was directly introduced. The operating conditions were as follows: reaction temperature 400°C; operating pressure 7.5 MPa; hydrogen-to-oil volume ratio 1500:1; volume space velocity 2.0 h -1 The heavy naphtha yield was 52.3% and the aromatic potential was 64.7.

[0054] Example 5

[0055] (1) The temperature of each bed of the stable operation wax oil hydrocracking unit was lowered to 190°C, and the feed was switched to direct current diesel. The wax oil hydrocracking reaction conditions were as follows: reaction temperature 355°C; operating pressure 15.2 MPa; hydrogen to oil volume ratio 1000:1; volume space velocity 1.5 h -1 .

[0056] (2) When the sulfur content of the hydrocracking unit reaction effluent is less than 10 μg·g -1 When the reaction conditions of the hydrocracking unit are adjusted so that the mass yield of the light oil of the fraction with the final boiling point of the reaction effluent <270°C reaches 4%, the temperature is then kept constant;

[0057] (3) During the constant temperature process of step (2), the hydrocracking catalyst in each bed was subjected to hydrogen reduction treatment. Fresh hydrogen was introduced into each bed along the logistics direction. The hydrogen-to-oil ratio of the fresh hydrogen introduced into the first catalyst was 200:1. The increase in the hydrogen introduced into the adjacent beds was 16%. The reduction time was 6 hours.

[0058] (4) After the reduction treatment, catalytic cracking diesel was directly introduced. The operating conditions were as follows: reaction temperature 405°C; operating pressure 9.3 MPa; hydrogen-to-oil volume ratio 1300:1; volume space velocity 2.5 h -1 The heavy naphtha yield was 54.8% and the aromatic potential was 66.2.

[0059] Example 6

[0060] (1) The temperature of each bed of the stable operation wax oil hydrocracking unit was lowered to 185°C, and the feed was switched to direct current diesel. The wax oil hydrocracking reaction conditions were as follows: reaction temperature 345°C; operating pressure 13.3 MPa; hydrogen to oil volume ratio 800:1; volume space velocity 2.5 h -1 .

[0061] (2) When the sulfur content of the hydrocracking unit reaction effluent is less than 150 μg·g -1 When the reaction conditions of the hydrocracking unit are adjusted so that the mass yield of the light oil of the fraction with the final boiling point of the reaction effluent < 220°C reaches 8%, the temperature is then kept constant;

[0062] (3) During the constant temperature process of step (2), the hydrocracking catalyst in each bed is subjected to hydrogen reduction treatment. Fresh hydrogen is introduced into each bed along the logistics direction. The hydrogen-to-oil ratio of fresh hydrogen introduced into the first catalyst is 800:1. The increase in the amount of hydrogen introduced into the adjacent beds is 25%, and the reduction time is 20 hours.

[0063] (4) After the reduction treatment, catalytic cracking diesel was directly introduced. The operating conditions were as follows: reaction temperature 415°C; operating pressure 6.3 MPa; hydrogen-to-oil volume ratio 1600:1; volume space velocity 1.7 h -1 The heavy naphtha yield was 56.2% and the aromatic potential was 65.3%.

Claims

1. A method for switching from wax oil hydrocracking to catalytic cracking diesel hydrocracking, characterized in that: The method includes the following contents: (1) The average temperature of each bed of the stable-operating wax oil hydrocracking unit is cooled to 90°C to 230°C, preferably 180°C to 230°C, and the feed is switched to direct-flow diesel; (2) When the sulfur content of the hydrocracking unit reaction effluent is less than 500 μg·g -1 When the concentration is less than 100 μg·g -1 , most preferably less than 10 μg·g -1 When the reaction conditions of the hydrocracking unit are adjusted, the mass yield of the light oil of the fraction with the final boiling point of the reaction effluent <280°C is 1% to 10%, and then the temperature is kept constant; (3) During the constant temperature process in step (2), the hydrocracking catalyst in each bed is subjected to reduction treatment, preferably with the reduction depth of the hydrocracking catalyst gradually increasing along the direction of the flow; (4) After the reduction treatment is completed, catalytic cracking diesel is directly introduced or catalytic cracking diesel is introduced after sulfurization to react to obtain naphtha with high yield and high aromatic potential.

2. The method according to claim 1, wherein: In step (1), the wax oil hydrocracking unit is provided with N hydrocracking catalyst beds, wherein N is at least 2, preferably 2 to 6, and most preferably 3 to 4; and the packing volume ratio of the Nth hydrocracking catalyst bed to the N-1th catalyst bed is 1:0.1 to 1:10, preferably 1:0.5 to 1:5.

0.

3. The method according to claim 1, wherein: In step (1), the wax oil feedstock is contacted with the hydrocracking catalyst during stable operation, and the effluent from the wax oil hydrocracking reaction is separated and fractionated to obtain wax oil hydrocracking products including light naphtha, heavy naphtha, aviation kerosene, diesel and tail oil.

4. The method according to claim 1, wherein: In step (1), the hydrocracking catalyst loaded into the wax oil hydrocracking unit can be a commercial product or prepared by conventional methods, or a regenerated hydrocracking catalyst can be used.

5. The method according to claim 1, wherein: In step (1), the cracking component of the hydrocracking catalyst loaded into the wax oil hydrocracking unit includes but is not limited to Y-type or β molecular sieve, and the hydrocracking catalyst includes a hydrogenation active metal, a molecular sieve component and an alumina carrier, and the hydrogenation active metal includes but is not limited to one or a mixture of Wo, Mo, Co, and Ni.

6. The method according to claim 1, wherein: In step (1), the hydrocracking catalyst comprises a carrier and a supported hydrogenation metal, and the hydrogenation metal comprises a Group VIB metal component and a Group VIII metal component of the periodic table based on the weight of the catalyst.

7. The method according to claim 1, wherein: In step (1), the wax oil is a vacuum wax oil, the final distillation point of the vacuum wax oil is 450-550° C., preferably 460-520° C., the density is below 0.92 g / cm 3 , preferably below 0.91 g / cm 3 , the sulfur content is above 0.1 wt %, preferably 0.1-3.0 wt %, and the aromatics content is below 60 wt %, preferably 1.0-50 wt %.

8. The method according to claim 1, wherein: In step (1), the reaction temperature of the wax oil hydrocracking is 250°C to 450°C, preferably 310°C to 390°C; the operating pressure is 2.0MPa to 20.0MPa, preferably 8MPa to 16.0MPa; the hydrogen-to-oil volume ratio is 100:1 to 2500:1, preferably 500:1 to 1500:1; and the volume space velocity is 0.1 to 5.0h -1 , preferably 0.5~3.0h -1 .

9. The method according to claim 1, wherein: In step (1), the straight-run diesel has an initial boiling point greater than 170°C, preferably 170-280°C; a final boiling point of 210-410°C, preferably 260-390°C; and a density of 0.65-0.90 g / cm 3 , preferably 0.70~0.85g / cm 3 ; Sulfur content is 500μg·g -1 Below, preferably 100 μg·g -1 the following.

10. The method according to claim 1, wherein: In step (2), the reaction conditions of the hydrocracking unit are adjusted so that the mass yield of light oil in the fraction with a final boiling point of less than 280° C. in the reaction effluent is 1% to 10%.

11. The method according to claim 1, wherein: In step (3), the reduction treatment is carried out by introducing hydrogen into each bed, and the reduction degree of the catalyst in different beds is controlled by adjusting and controlling the hydrogen injection time, hydrogen-to-oil ratio, hydrogen purity and H2S gas content in the hydrogen.

12. The method according to claim 1, wherein: In step (3), the hydrogen purity of the injected hydrogen is 30% to 100%, preferably 65% ​​to 100%, and the amount of H2S gas contained in the injected hydrogen is 0.1 to 1000 μl·L-1, preferably 1 to 100 μl·L-1.

13. The method according to claim 1, wherein: In step (3), the adjustment of the hydrogen injection time of each bed layer is to adjust the hydrogen injection time of the hydrogen injection equipment of each bed layer respectively. The hydrogen injection time of a single bed layer is 0.1h to 50h, preferably 1h to 20h.

14. The method according to claim 1, wherein: In step (3), the hydrogen-to-oil volume ratio of each bed is adjusted by adjusting the hydrogen injection amount of the hydrogen injection equipment of each bed to achieve the purpose of flexibly adjusting the hydrogen-to-oil ratio of each bed. The hydrogen-to-oil ratio of a single bed is 100:1 to 5000:1, preferably 300:1 to 2500:

1.

15. The method according to claim 1, wherein: In step (3), the reduction treatment method is to introduce fresh hydrogen into each bed along the logistics direction, with an increasing trend, the increase range of adjacent beds is 10%-30%, and the reduction time is 0.1h to 50h.

16. The method according to claim 1, wherein: In step (4), the catalytic cracking diesel has a final boiling point of 210 to 450° C., preferably 260 to 390° C., a density of 0.85 to 0.99 g / cm 3 , preferably 0.93 to 0.97 g / cm 3 , a sulfur content of more than 0.1 wt %, preferably 0.1 to 3.0 wt %, and an aromatics content of more than 60 wt %, preferably 80 to 90 wt %.

17. The method according to claim 1, wherein: In step (4), the catalytic cracking diesel hydrocracking reaction conditions are as follows: reaction temperature is 250°C to 450°C, preferably 330°C to 420°C; operating pressure is 2.0MPa to 20.0MPa, preferably 5.0MPa to 10.0MPa; hydrogen-to-oil volume ratio is 100:1 to 2500:1, preferably 500:1 to 1500:1; volume space velocity is 0.1 to 5.0h-1, preferably 0.5 to 2.0h-1.

Images