High flash jet fuel and method for producing high flash jet fuel

Abstract

The invention relates to the field of aviation fuel, and discloses a high-flash-point jet fuel and a production method thereof. The method comprises the following steps: mixing raw material oil with hydrogen, and feeding the mixture into a hydrofining reactor containing a hydrofining catalyst for hydrofining reaction to obtain a refined effluent; wherein the raw material oil is straight-run diesel oil, the density of the raw material oil at 20 DEG C is 820-840 kg / m < 3 >, and the sulfur content is not greater than 1.0 wt%; introducing the refined effluent into a hydrocracking reactor containing a hydrocracking catalyst, carrying out a hydrocracking reaction, and fractionating the obtained cracked effluent to obtain a naphtha fraction, a kerosene fraction, a diesel fraction I and a diesel fraction II; and carrying out gas-liquid separation on the diesel fraction I to respectively obtain a light component which can be circulated for fractionation and a heavy component which is taken as a high-flash-point jet fuel product and is discharged from the device. According to the method, the high-flash-point jet fuel can be stably produced at low cost, the flash point of the high-flash-point jet fuel is higher than 60 DEG C, and the quality of the high-flash-point jet fuel product is improved.

Description

Technical Field

[0001] The invention relates to the field of aviation fuel, in particular to high flash point jet fuel and a method for producing the high flash point jet fuel. Background Art

[0002] High flash point jet fuel is a special type of jet fuel with a high flash point value. Compared with ordinary jet fuel, it effectively improves the flash point of the fuel while maintaining other properties. Generally speaking, jet fuel refers to RP-3 fuel, which is mainly used in various civil aviation and military jet aircraft. However, for carrier-based aircraft on naval ships, the particularity of the combat environment determines that its oil use is inevitably affected by many unfavorable factors such as high temperature, high humidity, high salt environment and swaying of the ship. For example, it is required that the oil does not deteriorate and cannot spontaneously combust under high temperature conditions, especially that the fuel has good fuel safety under the condition of swaying of the ship. Therefore, higher requirements are put forward for carrier-based aircraft fuel than ordinary aviation kerosene. In order to adapt to the complex working environment of carrier-based aircraft, the fuel is required to have a higher flash point (generally not less than 60°C), good combustion, stability and low temperature properties. Therefore, high flash point jet fuel is mainly used as carrier-based aircraft fuel, and the current standard is GJB560A-97. The existing jet fuel production technology is mainly the hydrorefining technology and hydrocracking technology of straight run jet fuel fractions.

[0003] US4172815A discloses a single-stage cycle hydrocracking method for simultaneously producing jet fuel and diesel, and the process flow is as follows: after the heavy crude oil is hydrocracking, its product is fractionated to obtain a jet fuel fraction, a diesel fraction and tail oil, and the jet fuel fraction is mixed with the tail oil in whole or in part and sent back to the hydrocracking reactor. The obvious disadvantage of this process flow is that although the quality of jet fuel is improved by hydrocracking, its yield is greatly reduced, hydrogen consumption is increased, and investment is also increased.

[0004] CN1272524A discloses a process flow of medium-pressure hydrocracking and kerosene deep hydrotreating. The process is to carry out aromatic saturation of kerosene fraction with high aromatic content obtained in the medium-pressure hydrocracking process under the conditions of low pressure, high hydrogen purity and low reaction temperature, and the catalyst used is a catalyst containing Pt or Ni reduced metal. The process can well process kerosene fraction with high aromatic content to obtain qualified jet fuel, but the method requires more equipment and device investment, and the operation is more complicated.

[0005] At the same time, the existing technology for improving the flash point of jet fuel generally has defects such as complicated production process, harsh reaction conditions, large quality fluctuations, and a significant impact on the quality of upstream and downstream products, resulting in poor production stability and high costs. Summary of the invention

[0006] The purpose of the present invention is to overcome the problems in the prior art that the production method of high flash point jet fuel is high in cost and the flash point of the product is still low.

[0007] In order to achieve the above object, the first aspect of the present invention provides a method for producing high flash point jet fuel, the method comprising:

[0008] (1) The feedstock oil is mixed with hydrogen and then enters a hydrotreating reactor containing a hydrotreating catalyst to undergo a hydrotreating reaction to obtain a refined effluent;

[0009] The raw oil is straight-run diesel, and the density of the straight-run diesel at 20°C is 820-840 kg / m 3 , sulfur content ≯1.0wt%;

[0010] (2) introducing the refined effluent into a hydrocracking reactor containing a hydrocracking catalyst to carry out a hydrocracking reaction to obtain a cracking effluent;

[0011] (3) fractionating the cracking effluent to obtain a naphtha fraction, a kerosene fraction, a diesel fraction I and a diesel fraction II;

[0012] (4) separating the diesel fraction I into a gas-liquid phase to obtain a light component that can be recycled to step (3) for the fractionation and a heavy component that exits the device as a high flash point jet fuel product;

[0013] The high flash point jet fuel product has an initial distillation point of 175-195°C, a final distillation point of 245-260°C, and a flash point greater than 60°C.

[0014] The second aspect of the present invention provides a high flash point jet fuel prepared by the method described in the first aspect, wherein the high flash point jet fuel has an initial distillation point of 175-195°C, a final distillation point of 245-260°C, and a flash point greater than 60°C.

[0015] Through the above technical solution, the present invention has the following advantages:

[0016] (1) The method for producing high-flash-point jet fuel provided by the present invention can increase the flash point of the high-flash-point jet fuel produced to above 60° C. based on RP-3 jet fuel, thereby significantly improving the safety and reliability of the jet engine and meeting the special requirements of carrier-based aircraft applications;

[0017] (2) The method for producing high-flash-point jet fuel provided by the present invention not only improves the quality of high-flash-point jet fuel products, but also does not require special technical transformation of the existing production process, does not require additional investment, and only requires precise control of the operating conditions of each stage of the process, and the operating pressure and temperature are relatively mild; the entire preparation process has strong raw material adaptability, excellent production flexibility, simple preparation process, and can achieve low-cost mass production;

[0018] (3) Through the method for producing high flash point jet fuel provided by the present invention, in a preferred case, the high flash point jet fuel produced has a low sulfur content, thereby reducing environmental pollution.

[0019] In summary, the method for producing high flash point jet fuel provided by the present invention uses straight run diesel that meets specific performance parameter requirements as the main raw material (raw oil), combines a suitable process flow and catalyst system, regulates the course and depth of the reaction, and converts more components in the straight run diesel into a high flash point jet fuel product, whose flash point is greater than 60° C. The method can realize the stable production of high flash point jet fuel at low cost, and improves the quality of the high flash point jet fuel product. BRIEF DESCRIPTION OF THE DRAWINGS

[0020] Figure 1 It is a process flow diagram of a preferred specific implementation mode of the method of the present invention.

[0021] Description of Reference Numerals

[0022] 1. Raw oil 2. Hydrogen 3. Heating furnace

[0023] 4. Hydrofining reactor 5. Hydrocracking reactor 6. Medium temperature and high pressure separator

[0024] 7. Medium temperature low pressure separator 8. Cold high pressure separator 9. Cold low pressure separator

[0025] 10. Hydrogen sulfide stripping tower 11. First tower top reflux tank 12. Dry gas

[0026] 13. Second part light naphtha fraction 14. Product fractionation tower 15. Naphtha fraction

[0027] 16. Kerosene fraction 17. Diesel fraction I 18. Diesel fraction II

[0028] 19. Second tower top reflux tank 20. First part of heavy naphtha fraction

[0029] 21. The second part of heavy naphtha fraction

[0030] 22. Jet fuel stripping tower 23. Jet fuel products 24. Low freezing point diesel stripping tower

[0031] 25. High flash point jet fuel products 26. Circulating hydrogen desulfurization tower

[0032] 27. Part I Light Naphtha Fraction 28. Light Components DETAILED DESCRIPTION

[0033] The endpoints and any values ​​of the ranges disclosed in this article are not limited to the precise ranges or values, and these ranges or values ​​should be understood to include values ​​close to these ranges or values. For numerical ranges, the endpoint values ​​of each range, the endpoint values ​​of each range and the individual point values, and the individual point values ​​can be combined with each other to obtain one or more new numerical ranges, which should be considered as specifically disclosed in this article.

[0034] As mentioned above, the first aspect of the present invention provides a method for producing high flash point jet fuel, the method comprising:

[0035] (1) The feedstock oil is mixed with hydrogen and then enters a hydrotreating reactor containing a hydrotreating catalyst to undergo a hydrotreating reaction to obtain a refined effluent;

[0036] The raw oil is straight-run diesel, and the density of the straight-run diesel at 20°C is 820-840 kg / m 3 , sulfur content ≯1.0wt%;

[0037] (2) introducing the refined effluent into a hydrocracking reactor containing a hydrocracking catalyst to carry out a hydrocracking reaction to obtain a cracking effluent;

[0038] (3) fractionating the cracking effluent to obtain a naphtha fraction, a kerosene fraction, a diesel fraction I and a diesel fraction II;

[0039] (4) separating the diesel fraction I into a gas-liquid phase to obtain a light component that can be recycled to step (3) for the fractionation and a heavy component that exits the device as a high flash point jet fuel product;

[0040] The high flash point jet fuel product has an initial distillation point of 175-195°C, a final distillation point of 245-260°C, and a flash point greater than 60°C.

[0041] The inventor of the present invention found in the research process that straight-run diesel that meets the requirements of specific performance parameters is first mixed with hydrogen as a raw oil, and then subjected to hydrofining reaction and hydrocracking reaction in sequence, and then the obtained cracking effluent is fractionated to obtain naphtha fraction, kerosene fraction, diesel fraction I and diesel fraction II; the diesel fraction I is subjected to gas-liquid separation to obtain light components that can be recycled to step (3) for the fractionation and heavy components that can be used as high flash point jet fuel products to exit the device, and its properties can meet the main indicators of GJB560A-97 high flash point jet fuel. The method can realize the production of high flash point jet fuel products without the need to be stopped at all times, while still ensuring the quality of other upstream and downstream products. Among them, the naphtha fraction can be used as a reforming raw material or a cracking raw material, the kerosene fraction can be used to produce aviation kerosene products (RP-3 jet fuel), and the diesel fraction II can meet the index requirements of Beijing Standard VI diesel.

[0042] According to some embodiments of the present invention, preferably, in step (1), the straight-run diesel has an initial distillation point of 187-213°C and a final distillation point of 322-349°C.

[0043] According to some embodiments of the present invention, preferably, the density of the straight-run diesel at 20°C is 822-830 kg / m 3 ; The sulfur content is 0.4-0.85wt%.

[0044] According to some embodiments of the present invention, preferably, the nitrogen content of the straight-run diesel is ≯150 mg / kg, preferably 0-120 mg / kg.

[0045] According to some embodiments of the present invention, preferably, the carbon residue value of the straight-run diesel is less than 0.1 wt %.

[0046] The use of straight-run diesel that meets the above-mentioned specific parameter requirements is conducive to obtaining a high-flash point jet fuel product with a higher flash point and better quality.

[0047] According to some embodiments of the present invention, preferably, in step (1), the hydrotreating catalyst comprises a carrier and an active component supported on the carrier; the carrier is alumina; the active component is selected from at least two of Co, Mo, Ni and W, preferably Mo and Ni;

[0048] Preferably, based on the total weight of the hydrotreating catalyst, the content of Mo is 22-25wt%, the content of Ni is 3.6-4.2wt% in terms of oxides; and the content of the carrier is 68-72wt%.

[0049] The use of the above preferred embodiment is beneficial to the removal of impurities such as sulfur, nitrogen, oxygen, metals, and the hydrogenation saturation of unsaturated hydrocarbons.

[0050] According to some embodiments of the present invention, preferably, in step (1), the reaction conditions of the hydrofining reaction include: a reaction temperature of 300-400°C, a reaction pressure of 6-10 MPa, a hydrogen-oil volume ratio of (500-1000): 1, and a volume space velocity of 2-5 h -1 .

[0051] According to some embodiments of the present invention, preferably, before performing the hydrotreating reaction, step (1) further comprises: preheating the feedstock oil and hydrogen to 350-400° C. respectively, and then performing the mixing.

[0052] According to some embodiments of the present invention, preferably, in step (2), the reaction conditions of the hydrocracking reaction include: a reaction temperature of 300-400°C, a reaction pressure of 5-10 MPa, a hydrogen-to-oil volume ratio of (500-1000): 1, and a volume space velocity of 1-4 h -1 .

[0053] According to some embodiments of the present invention, preferably, in step (2), the hydrocracking catalyst comprises a carrier and an active component supported on the carrier; the carrier is alumina; the active component is selected from at least two of Co, Mo, Ni and W, preferably Ni and W;

[0054] Preferably, based on the total weight of the hydrocracking catalyst, the content of W is 22-25wt%, the content of Ni is 6.0-7.5wt% in terms of oxides; and the content of the carrier is 70-74wt%.

[0055] The above preferred embodiment is advantageous for breaking the chains of macromolecular hydrocarbons into small molecule hydrocarbons.

[0056] According to some embodiments of the present invention, preferably, in step (3), the fractionation is carried out in a fractionation tower, and before the fractionation is carried out, the method further comprises: according to the flow direction of the reactant stream, a high-pressure separator and a low-pressure separator are arranged between the hydrocracking reactor and the fractionation tower to separate the cracking effluent into gas and liquid, and then the obtained liquid product is introduced into a hydrogen sulfide stripping tower to remove the hydrogen sulfide therein.

[0057] According to some embodiments of the present invention, preferably, in step (3), the top temperature of the fractionation tower is 110-125°C, and the bottom temperature is 260-270°C.

[0058] According to some embodiments of the present invention, preferably, in step (3), the extraction temperature of the kerosene fraction is 150-158°C; and the extraction temperature of the diesel fraction I is 180-185°C.

[0059] According to some embodiments of the present invention, preferably, in step (3), the extraction amount of the kerosene fraction is 5-20 t / h.

[0060] According to some embodiments of the present invention, preferably, the method further comprises: separating the kerosene fraction into gas and liquid to obtain a gas phase which can be recycled to step (3) for the fractionation and a liquid phase which can be discharged from the device as a jet fuel product.

[0061] According to some embodiments of the present invention, preferably, the gas-liquid separation is carried out in a jet fuel stripping tower, and the conditions for the gas-liquid separation include: the bottom temperature of the jet fuel stripping tower is 170-200°C.

[0062] According to some embodiments of the present invention, preferably, the distillation range of the liquid phase used as the jet fuel product outlet device is 160-233°C.

[0063] According to some embodiments of the present invention, preferably, the density of the liquid phase at 20°C as the jet fuel product outlet device is 792-797 kg / m 3 .

[0064] According to some embodiments of the present invention, preferably, the freezing point of the liquid phase of the jet fuel product outlet device is less than -60°C.

[0065] According to some embodiments of the present invention, preferably, the flash point of the liquid phase of the jet fuel product outlet device is 43-50°C.

[0066] According to some embodiments of the present invention, preferably, the smoke point of the liquid phase of the jet fuel product outlet device is 21-23 mm.

[0067] According to some embodiments of the present invention, preferably, in step (4), the gas-liquid separation is carried out in a low-freezing diesel stripper, and the conditions for the gas-liquid separation include: the bottom temperature of the low-freezing diesel stripper is 200-225° C.; the extraction rate of the high flash point jet fuel product is 10-18 t / h; and the return tower temperature is 210-230° C. The use of the above preferred embodiment is conducive to further obtaining a high flash point jet fuel product with better quality.

[0068] The second aspect of the present invention provides a high flash point jet fuel prepared by the method described in the first aspect, wherein the high flash point jet fuel has an initial distillation point of 175-195°C, a final distillation point of 245-260°C, and a flash point greater than 60°C.

[0069] According to some embodiments of the present invention, preferably, the flash point of the high flash point jet fuel is 61-75°C.

[0070] According to some embodiments of the present invention, preferably, the total sulfur content of the high flash point jet fuel is ≯0.005 wt %.

[0071] According to some embodiments of the present invention, preferably, the smoke point of the high flash point jet fuel is greater than 20 mm, more preferably 20.5-22.5 mm.

[0072] According to some embodiments of the present invention, preferably, the high flash point jet fuel has a density of 804-810 kg / m at 20°C. 3 .

[0073] According to some embodiments of the present invention, preferably, the freezing point of the high flash point jet fuel is ≯-47°C.

[0074] According to some embodiments of the present invention, appropriate amounts of additives such as aviation kerosene antioxidants, aviation kerosene antistatic agents and aviation kerosene antiwear agents can be added to ensure that the high flash point jet fuel product has good oxidation stability, lubricity and safety.

[0075] The method provided by the present invention is further described in detail below with reference to the accompanying drawings.

[0076] like Figure 1 As shown, in a preferred specific embodiment of the method of the present invention, the raw oil (straight-run diesel) 1 is first heated by a heat exchanger, and the hydrogen 2 is heated by a heating furnace 3, and then the heated raw oil 1 and the heated hydrogen 2 are mixed and introduced into a hydrorefining reactor 4 containing a hydrorefining catalyst for hydrorefining reaction to obtain a refined effluent;

[0077] The refined effluent is sent to a hydrocracking reactor 5 containing a hydrocracking catalyst to undergo a hydrocracking reaction to obtain a cracking effluent;

[0078] After the cracking effluent is heat exchanged with the feedstock oil 1 through a heat exchanger, it is sent to a medium-temperature high-pressure separator 6 for hot high-pressure gas-liquid separation to obtain a gas phase product I and a liquid phase product I; the liquid phase product I is sent to a medium-temperature low-pressure separator 7 for hot low-pressure gas-liquid separation to obtain a gas phase product II and a liquid phase product II; the gas phase product I is sent to a cold high-pressure separator 8 for gas-liquid separation to obtain a gas phase product III and a liquid phase product III;

[0079] The gas phase product III is sent to the circulating hydrogen desulfurization tower 26 to remove hydrogen sulfide and then recycled;

[0080] After the liquid phase product III and the gas phase product II obtained by the medium temperature low pressure separator 7 are sent to the cold low pressure separator 9 for separation, the liquid phase product IV at the bottom of the cold low pressure separator 9 and the liquid phase product II are sent to the hydrogen sulfide stripping tower 10 to remove the hydrogen sulfide therein, and the gas phase at the top of the tower is sent to the first tower top reflux tank 11 for three-phase separation of oil phase, gas phase and water phase to obtain water phase, dry gas 12 and light naphtha fraction, wherein the light naphtha fraction is divided into a first part of light naphtha fraction 27 and a second part of light naphtha fraction 13, the first part of light naphtha fraction 27 is returned to the hydrogen sulfide stripping tower 10, and the second part of light naphtha fraction 13 is discharged from the device;

[0081] Preferably, the volume flow ratio of the first portion of light naphtha fraction to the second portion of light naphtha fraction is (4-5):1;

[0082] The bottom liquid phase of the hydrogen sulfide stripping tower 10 is sent to the product fractionation tower 14 for fractionation and cutting to obtain a naphtha fraction 15, a kerosene fraction 16, a diesel fraction I 17 and a diesel fraction II 18; wherein:

[0083] The naphtha fraction 15 is sent to the second tower top reflux tank 19 for gas-liquid separation to obtain a light component 28 and a heavy naphtha fraction, wherein the heavy naphtha fraction is divided into a first part of a heavy naphtha fraction 20 and a second part of a heavy naphtha fraction 21, the first part of the heavy naphtha fraction 20 is returned to the product fractionation tower 14, and the second part of the heavy naphtha fraction 21 is discharged from the device;

[0084] Preferably, the volume flow ratio of the first portion of heavy naphtha fraction to the second portion of heavy naphtha fraction is (1-2): 1;

[0085] The kerosene fraction 16 enters the aviation kerosene stripping tower 22 for gas-liquid separation, the light components at the top of the tower are returned to the product fractionation tower 14, and the liquid phase at the bottom of the tower is discharged from the device as the aviation kerosene product 23;

[0086] The diesel fraction I17 enters the low-freezing point diesel stripping tower 24 for gas-liquid separation, the light components at the top of the tower are returned to the product fractionation tower 14, and the heavy components at the bottom of the tower are discharged from the device as a high flash point jet fuel product 25.

[0087] The present invention will be described in detail below through examples.

[0088] In the following examples and comparative examples, unless otherwise specified, all raw materials used are commercially available. The properties of the raw material oil used are shown in Table 1;

[0089] The composition of the hydrotreating catalyst used is: the carrier is alumina; the active component is MoO 3, NiO; based on the total weight of the hydrorefining catalyst, MoO 3 The content of is 25wt%, the content of NiO is 4.2wt%, and the balance is carrier;

[0090] The hydrocracking catalyst used is composed of: alumina as carrier; NiO and WO as active components 3 Based on the total weight of the hydrotreating catalyst, WO 3 The content of is 22wt%, the content of NiO is 6.0wt%, and the balance is carrier.

[0091] Table 1: Raw oil properties

[0092] Property parameters unit Test Method Crude Oil A Crude oil B Machine Miscellaneous / Visual inspection none none Density at 20℃ <![CDATA[kg / m 3 ]]> SH / T0604 827 824 Nitrogen content mg / kg SH / T 0657 106 74 Sulfur content wt% GB / T 17040 0.77 0.73 Chlorine content mg / kg Q / SH3155.S01.078 ＜0.5 ＜0.5 Initial distillation point ℃ GB / T 6536 198 190 Final distillation point ℃ GB / T 6536 334 330 Moisture / Visual inspection none none Carbon residue wt% GB / T17144 ＜0.1 ＜0.1 Total metal content μg / g Q / SH3155.S01.500 ＜1 ＜1 C7 insoluble matter μg / g Q / SH3155.SO1.130 ＜100 ＜100

[0093] Examples 1-6 are used to illustrate the method provided by the present invention

[0094] Example 1

[0095] Use Figure 1 The process flow shown is carried out, wherein:

[0096] The straight-run diesel oil used is the raw material oil A; before the hydrofining reaction, the raw material oil and hydrogen are preheated to 350°C respectively and then mixed;

[0097] The reaction conditions of the hydrotreating reaction are: reaction temperature of 355°C, reaction pressure of 8.5MPa, hydrogen-oil volume ratio of 700:1, volume space velocity of 3h -1 ;

[0098] The reaction conditions of the hydrocracking reaction are: reaction temperature of 371°C, reaction pressure of 8.5MPa, hydrogen-oil volume ratio of 700:1, volume space velocity of 2h -1 ;

[0099] The top temperature of the fractionation tower is 120°C and the bottom temperature is 268°C;

[0100] The extraction temperature of the kerosene fraction is 152°C; the extraction temperature of the diesel fraction I is 183°C;

[0101] The extraction rate of kerosene fraction is 8t / h; the conditions for gas-liquid separation of kerosene fraction in the aviation kerosene stripper are: the bottom temperature of the aviation kerosene stripper is 182℃;

[0102] The conditions for gas-liquid separation of diesel fraction I in the low-freezing diesel stripper are as follows: the bottom temperature of the low-freezing diesel stripper is 210°C; the extraction rate of high flash point jet fuel product is 14.5t / h; the return tower temperature is 214°C;

[0103] The volume flow ratio of the first part of light naphtha fraction to the second part of light naphtha fraction is 4.9:1; the volume flow ratio of the first part of heavy naphtha fraction to the second part of heavy naphtha fraction is 1.5:1.

[0104] Example 2

[0105] The method of Example 1 is different from that of Example 1:

[0106] The straight-run diesel used is raw material oil B;

[0107] The reaction conditions of the hydrotreating reaction are: reaction temperature of 355°C, reaction pressure of 8.5MPa, hydrogen-oil volume ratio of 700:1, and volume space velocity of 3h -1 ;

[0108] The reaction conditions of the hydrocracking reaction are: reaction temperature of 371°C, reaction pressure of 8.5MPa, hydrogen-oil volume ratio of 700:1, volume space velocity of 2h -1 ;

[0109] The top temperature of the fractionation tower is 122°C and the bottom temperature is 265°C;

[0110] The extraction temperature of the kerosene fraction is 155°C, and the extraction temperature of the diesel fraction I is 183°C;

[0111] The extraction rate of kerosene fraction is 10t / h; the conditions for gas-liquid separation of kerosene fraction in the aviation kerosene stripper are: the bottom temperature of the aviation kerosene stripper is 184℃;

[0112] The conditions for gas-liquid separation of diesel fraction I in the low-freezing point diesel stripping tower are: the bottom temperature of the low-freezing point diesel stripping tower is 210°C; the extraction rate of high flash point jet fuel product is 14.5t / h; the return tower temperature is 214°C; and the rest are the same.

[0113] Example 3

[0114] The method of Example 1 is different from that of Example 1:

[0115] The reaction conditions of the hydrotreating reaction are: reaction temperature of 355°C, reaction pressure of 8.5MPa, hydrogen-oil volume ratio of 700:1, and volume space velocity of 3h -1 ;

[0116] The reaction conditions of the hydrocracking reaction are: reaction temperature of 371°C, reaction pressure of 8.5MPa, hydrogen-oil volume ratio of 700:1, volume space velocity of 2h -1 ;

[0117] The top temperature of the fractionation tower is 124°C and the bottom temperature is 265°C;

[0118] The extraction temperature of the kerosene fraction is 155°C, and the extraction temperature of the diesel fraction I is 183°C;

[0119] The extraction rate of kerosene fraction is 10t / h; the conditions for gas-liquid separation of kerosene fraction in the aviation kerosene stripper are: the bottom temperature of the aviation kerosene stripper is 184℃;

[0120] The conditions for gas-liquid separation of diesel fraction I in the low-freezing point diesel stripping tower are: the bottom temperature of the low-freezing point diesel stripping tower is 210°C; the extraction rate of high flash point jet fuel product is 14.5t / h; the return tower temperature is 214°C; and the rest are the same.

[0121] Example 4

[0122] The method of Example 1 is different from that of Example 1:

[0123] The reaction conditions of the hydrotreating reaction are: reaction temperature of 355°C, reaction pressure of 8.5MPa, hydrogen-oil volume ratio of 700:1, volume space velocity of 3h -1 ;

[0124] The reaction conditions of the hydrocracking reaction are: reaction temperature of 371°C, reaction pressure of 8.5MPa, hydrogen-oil volume ratio of 700:1, volume space velocity of 2h -1 ;

[0125] The top temperature of the fractionation tower is 120°C and the bottom temperature is 268°C;

[0126] The extraction temperature of the kerosene fraction is 152°C, and the extraction temperature of the diesel fraction I is 180°C;

[0127] The extraction rate of kerosene fraction is 8t / h; the conditions for gas-liquid separation of kerosene fraction in the aviation kerosene stripper are: the bottom temperature of the aviation kerosene stripper is 180℃;

[0128] The conditions for gas-liquid separation of diesel fraction I in the low-freezing diesel stripping tower are: the bottom temperature of the low-freezing diesel stripping tower is 207°C; the extraction rate of high flash point jet fuel product is 13.5t / h; the return tower temperature is 214°C; and the rest are the same.

[0129] Example 5

[0130] The method of Example 1 is different from that of Example 1:

[0131] The reaction conditions of the hydrotreating reaction are: reaction temperature of 355°C, reaction pressure of 8.5MPa, hydrogen-oil volume ratio of 700:1, volume space velocity of 3h -1 ;

[0132] The reaction conditions of the hydrocracking reaction are: reaction temperature of 371°C, reaction pressure of 8.5MPa, hydrogen-oil volume ratio of 700:1, volume space velocity of 2h -1 ;

[0133] The top temperature of the fractionation tower is 122°C and the bottom temperature is 265°C;

[0134] The extraction temperature of the kerosene fraction is 155°C, and the extraction temperature of the diesel fraction I is 180°C;

[0135] The extraction rate of kerosene fraction is 10t / h; the conditions for gas-liquid separation of kerosene fraction in the aviation kerosene stripper are: the bottom temperature of the aviation kerosene stripper is 180℃;

[0136] The conditions for gas-liquid separation of diesel fraction I in the low-freezing diesel stripping tower are: the bottom temperature of the low-freezing diesel stripping tower is 207°C; the extraction rate of high flash point jet fuel product is 13t / h; the return tower temperature is 214°C; and the rest are the same.

[0137] Example 6

[0138] The method of Example 1 is different from that of Example 1:

[0139] The reaction conditions of the hydrotreating reaction are: reaction temperature of 355°C, reaction pressure of 8.5MPa, hydrogen-oil volume ratio of 700:1, volume space velocity of 3h -1 ;

[0140] The reaction conditions of the hydrocracking reaction are: reaction temperature of 371°C, reaction pressure of 8.5MPa, hydrogen-oil volume ratio of 700:1, volume space velocity of 2h -1 ;

[0141] The top temperature of the fractionation tower is 118°C and the bottom temperature is 265°C;

[0142] The extraction temperature of the kerosene fraction is 150°C, and the extraction temperature of the diesel fraction I is 183°C;

[0143] The extraction rate of kerosene fraction is 7t / h; the conditions for gas-liquid separation of kerosene fraction in the aviation kerosene stripper are: the bottom temperature of the aviation kerosene stripper is 178℃;

[0144] The conditions for gas-liquid separation of diesel fraction I in the low-freezing point diesel stripping tower are: the bottom temperature of the low-freezing point diesel stripping tower is 210°C; the extraction rate of high flash point jet fuel product is 14.5t / h; the return tower temperature is 214°C; and the rest are the same.

[0145] The yield and properties of the products obtained in each embodiment are shown in Table 2.

[0146] Table 2

[0147]

[0148]

[0149] It can be seen from the above results that the method provided by the present invention can realize the stable production of high flash point jet fuel with a low cost, the flash point of which is greater than 60° C., and improves the quality of the high flash point jet fuel product.

[0150] The preferred embodiments of the present invention are described in detail above, but the present invention is not limited thereto. Within the technical concept of the present invention, the technical solution of the present invention can be subjected to a variety of simple modifications, including the combination of various technical features in any other suitable manner, and these simple modifications and combinations should also be regarded as the contents disclosed by the present invention and belong to the protection scope of the present invention.

Claims

1. A method for producing high flash point jet fuel, characterized in that: The method comprises: (1) The feedstock oil is mixed with hydrogen and then enters a hydrotreating reactor containing a hydrotreating catalyst to undergo a hydrotreating reaction to obtain a refined effluent; The raw oil is straight-run diesel, and the density of the straight-run diesel at 20°C is 820-840 kg / m 3 , sulfur content ≯1.0wt%; (2) introducing the refined effluent into a hydrocracking reactor containing a hydrocracking catalyst to carry out a hydrocracking reaction to obtain a cracking effluent; (3) fractionating the cracking effluent to obtain a naphtha fraction, a kerosene fraction, a diesel fraction I and a diesel fraction II; (4) separating the diesel fraction I into a gas-liquid phase to obtain a light component that can be recycled to step (3) for the fractionation and a heavy component that exits the device as a high flash point jet fuel product; The high flash point jet fuel product has an initial distillation point of 175-195°C, a final distillation point of 245-260°C, and a flash point greater than 60°C.

2. The method according to claim 1, wherein: In step (1), the straight-run diesel has an initial boiling point of 187-213°C and a final boiling point of 322-349°C; Preferably, the density of the straight-run diesel at 20°C is 822-830 kg / m 3 ; Sulfur content is 0.4-0.85wt%; Nitrogen content ≯150mg / kg, preferably 0-120mg / kg; The residual carbon value is less than 0.1wt%.

3. The method according to claim 1 or 2, wherein: In step (1), the hydrotreating catalyst comprises a carrier and an active component supported on the carrier; the carrier is alumina; the active component is selected from at least two of Co, Mo, Ni and W, preferably Mo and Ni; Preferably, based on the total weight of the hydrotreating catalyst, the content of Mo is 22-25wt%, the content of Ni is 3.6-4.2wt% in terms of oxides; and the content of the carrier is 68-72wt%.

4. The method according to any one of claims 1 to 3, wherein: In step (1), the reaction conditions of the hydrotreating reaction include: reaction temperature of 300-400°C, reaction pressure of 6-10 MPa, hydrogen-oil volume ratio of (500-1000):1, volume space velocity of 2-5 h -1 ; Preferably, before the hydrofining reaction is carried out, step (1) further comprises: preheating the feedstock oil and hydrogen to 350-400° C. respectively, and then carrying out the mixing.

5. The method according to any one of claims 1 to 4, wherein: In step (2), the reaction conditions of the hydrocracking reaction include: reaction temperature of 300-400°C, reaction pressure of 5-10 MPa, hydrogen-oil volume ratio of (500-1000):1, volume space velocity of 1-4 h -1 .

6. The method according to any one of claims 1 to 5, wherein: In step (2), the hydrocracking catalyst comprises a carrier and an active component supported on the carrier; the carrier is alumina; the active component is selected from at least two of Co, Mo, Ni and W, preferably Ni and W; Preferably, based on the total weight of the hydrocracking catalyst, the content of W is 22-25wt%, the content of Ni is 6.0-7.5wt% in terms of oxides; and the content of the carrier is 70-74wt%.

7. The method according to any one of claims 1 to 6, wherein: In step (3), the fractionation is carried out in a fractionation tower. Before the fractionation is carried out, the method further comprises: arranging a high-pressure separator and a low-pressure separator between the hydrocracking reactor and the fractionation tower according to the flow direction of the reactant stream to separate the cracking effluent into gas and liquid, and then introducing the obtained liquid product into a hydrogen sulfide stripping tower to remove hydrogen sulfide therein; Preferably, the top temperature of the fractionation tower is 110-125°C, and the bottom temperature is 260-270°C. Preferably, in step (3), the extraction temperature of the kerosene fraction is 150-158°C; the extraction temperature of the diesel fraction I is 180-185°C; Preferably, in step (3), the extraction rate of the kerosene fraction is 5-20 t / h.

8. The method according to any one of claims 1 to 7, wherein: The method further comprises: separating the kerosene fraction into gas and liquid to obtain a gas phase that can be recycled to step (3) for fractionation and a liquid phase that exits the device as a jet fuel product; Preferably, the gas-liquid separation is carried out in a jet fuel stripping tower, and the conditions for the gas-liquid separation include: the bottom temperature of the jet fuel stripping tower is 170-200°C; Preferably, the liquid phase of the jet fuel product outlet device has a distillation range of 160-233°C and a density of 792-797 kg / m 3 ; Freezing point is less than -60℃; flash point is 43-50℃; smoke point is 21-23mm.

9. The method according to any one of claims 1 to 8, wherein: In step (4), the gas-liquid separation is carried out in a low-freezing-point diesel stripping tower, and the conditions for the gas-liquid separation include: the bottom temperature of the low-freezing-point diesel stripping tower is 200-225°C; the extraction rate of the high flash point jet fuel product is 10-18t / h; and the return tower temperature is 210-230°C.

10. A high flash point jet fuel prepared by the method according to any one of claims 1 to 9, characterized in that: The high flash point jet fuel has an initial boiling point of 175-195°C, a final boiling point of 245-260°C, and a flash point greater than 60°C; Preferably, the high flash point jet fuel has a flash point of 61-75°C; a total sulfur content of ≯0.005wt%; Preferably, the smoke point of the high flash point jet fuel is greater than 20 mm, preferably 20.5-22.5 mm; the density at 20°C is 804-810 kg / m 3 ; Freezing point ≯-47℃.

Images