A method for producing jet fuel by hydrogenation

Abstract

The present invention discloses a method for producing jet fuel by hydrogenation. The method comprises the following steps: (1) The jet fuel raw material and hydrogen enter a first reactor to carry out gas-phase hydrogenation reaction to obtain an effluent from the first reactor; (2) The effluent from the first reactor enters a second reactor after being pressurized. Among them, the gas-phase components are discharged upward from the second reactor, and the liquid-phase components undergo liquid-phase hydrogenation reaction downward to obtain hydrogenation heavy components; (3) The gas-phase components discharged from the second reactor are separated to obtain hydrogenation light components and hydrogen containing hydrogen sulfide; (4) After the hydrogenation heavy components and the hydrogenation light components are mixed, they are separated to obtain jet fuel products. The present invention can produce high-quality jet fuel products with high smoke points under the conditions of simple process, low energy consumption and low cost.

Description

Technical Field

[0001] The present invention belongs to the field of clean oil refining, and particularly relates to a method for hydrogenating and producing aviation kerosene. Background Art

[0002] In recent years, with the continuous development of China's civil aviation industry, the demand for aviation fuel has been increasing continuously. On the other hand, with the adjustment of China's energy structure, the demand for diesel has been slowing down year by year, and reducing the diesel-to-gasoline ratio has become the overall development trend of the refinery structure adjustment. Increasing the cut-off point of the aviation kerosene fraction, that is, increasing the final boiling point of the aviation kerosene and cutting more diesel components into the aviation kerosene, can significantly increase the yield of aviation kerosene and reduce the output of diesel, thus well solving the problem of the sharp increase in the current demand for aviation kerosene and the decrease in the demand for diesel. However, as the aviation kerosene fraction becomes heavier, some aromatics in the diesel return to the aviation kerosene component, and the increase in the aromatics content in the aviation kerosene will lead to the unqualified situation of the smoke point index of the aviation kerosene product. The pressure design level of the existing aviation kerosene unit is relatively low, generally 3 - 4 MPa, while aromatics hydrogenation requires a relatively high reaction pressure to be achieved under the reaction environment of a fixed bed, and it is difficult for the existing unit to use a pressurization means to remove aromatics, so there is no significant improvement in the smoke point index. Therefore, although the limit of the final boiling point in the aviation kerosene quality index is ≯300 °C, in order to ensure the qualification of the smoke point index, refinery enterprises can only control the final boiling point of the aviation kerosene at ≯260 °C at present, and the purpose of increasing the production of aviation kerosene and reducing the production of diesel cannot be achieved.

[0003] CN107233927A discloses a medium-oil type hydrocracking catalyst and its preparation method. The β zeolite obtained by modification in this method has relatively balanced cracking performance, and the produced middle distillate oil has the characteristic of high smoke point of aviation kerosene. However, the catalyst used for producing aviation kerosene adopts a conventional fixed-bed hydrogenation process technology, which has the disadvantages of high hydrogen consumption and energy consumption and low yield of aviation kerosene.

[0004] CN111088072A discloses a hydrocracking method for reducing the bromine index of heavy naphtha and increasing the smoke point of aviation kerosene. This method is to recycle the aviation kerosene light fraction rich in aromatics obtained by fractionating the oil products after hydrofining and hydrocracking back to the hydrocracking unit, while the fraction rich in paraffins obtained by fractionation is used as the aviation kerosene product. However, the yield of the aviation kerosene produced by this method is relatively low, the energy consumption of the unit is large during the production process, and the product has a high content of paraffins, which is likely to cause the product to be unqualified. Summary of the Invention

[0005] Aiming at the deficiencies of the prior art, the present invention provides a method for hydrogenating and producing aviation kerosene. The present invention can produce high-quality aviation kerosene products with high smoke point under the conditions of simple process, low energy consumption and low cost.

[0006] The present invention provides a method for hydrogenating and producing aviation kerosene, which comprises the following steps:

[0007] (1) The jet fuel feedstock and hydrogen enter the first reactor for gas-phase hydrogenation reaction to obtain the effluent from the first reactor.

[0008] (2) The effluent from the first reactor enters the second reactor after being pressurized. Among them, the gas-phase components are discharged upward from the second reactor, and the liquid-phase components undergo liquid-phase hydrogenation reaction downward to obtain hydrogenated heavy components.

[0009] (3) The gas-phase components discharged from the second reactor are separated to obtain hydrogenated light components and hydrogen containing hydrogen sulfide.

[0010] (4) After the hydrogenated heavy components and the hydrogenated light components are mixed, they are separated to obtain jet fuel products.

[0011] Furthermore, the properties of the jet fuel feedstock are as follows: the initial boiling point is 100°C to 150°C, the final boiling point is 280°C to 320°C, the smoke point is not less than 10 mm, further 15 to 22 mm, the S content is 600 to 3000 μg / g, and the N content is 5 to 300 μg / g.

[0012] Furthermore, both the first reactor and the second reactor are fixed-bed reactors. The second reactor preferably adopts a fixed-bed reactor provided with a flash zone.

[0013] Furthermore, a flash zone is provided in the second reactor. No catalyst is filled above and in the flash zone, and the area below the flash zone is the reaction zone. The effluent from the first reactor is pressurized and fed into the flash zone of the second reactor. The obtained gas-phase components are discharged upward from the second reactor, and the obtained liquid-phase components undergo liquid-phase hydrogenation reaction downward. The obtained product, i.e., the hydrogenated heavy components, is discharged from the bottom of the second reactor.

[0014] Furthermore, the first reactor is a gas-phase hydrogenation reactor. The operating conditions of the first reactor are as follows: the pressure is 0.1 to 3.0 MPa, preferably 0.5 to 2.0 MPa; the hydrogen-oil volume ratio is 100 to 1000, preferably 200 to 600; the temperature is 150 to 300°C, preferably 200 to 280°C; the volume space velocity is 0.1 to 6.0 h -1 , preferably 1.0 to 4.0 h -1 .

[0015] Further, the catalyst loaded in the first reactor includes a hydrofining catalyst. The hydrofining catalyst includes a carrier and a hydrogenation active metal; wherein the carrier is an inorganic refractory oxide, selected from one or more of alumina, amorphous silica-alumina, silica, or titanium oxide, etc.; the hydrogenation active metal includes Group VIB and / or Group VIII metal components, wherein Group VIB is selected from tungsten and / or molybdenum, and Group VIII is selected from nickel and / or cobalt. Generally, based on the mass of the catalyst, the content of the Group VIB metal in terms of oxide is 10 wt% to 20 wt%, and the content of the Group VIII metal in terms of oxide is 1 wt% to 9 wt%. The hydrofining catalyst can be a commercial hydrofining catalyst, such as the FH-40 series catalysts developed by Fushun Research Institute of Petroleum and Petrochemicals, SINOPEC, for example, FH-40A and FH-40C catalysts.

[0016] Further, the effluent from the first reactor is pressurized by a compressor, and the compressor can be a conventional commercial compressor, such as a reciprocating or centrifugal compressor. Among them, the pressurization can ensure the normal feeding of the second reactor and meet the operating pressure requirements of the second reactor.

[0017] Further, the operating conditions of the second reactor are as follows: the pressure is 1.0 to 6.0 MPa, preferably 1.5 to 5.0 MPa, and further preferably 2.5 to 5.0 MPa; the temperature is 100 to 230 °C, preferably 100 to 160 °C; the volume space velocity is 0.5 to 4.0 h -1 , preferably 1.0 to 3.0 h -1 .

[0018] Further, the pressure of the second reactor is at least 0.5 MPa higher than the pressure of the first reactor, preferably 1.0 to 6.0 MPa higher, and further preferably 2.0 to 4.0 MPa higher.

[0019] Furthermore, the catalyst filled in the second reactor is a hydrodearomatization catalyst, which includes a support and a hydrogenation active metal. The support is an inorganic refractory oxide, generally selected from one or more of alumina, amorphous silica-alumina, silica, titanium oxide, etc., preferably alumina. The hydrogenation active metal includes Group VIB and / or Group VIII metal components, which can be noble metals or non-noble metals. The Group VIB metal is selected from tungsten and / or molybdenum, the Group VIII non-noble metal is selected from nickel and / or cobalt, and the Group VIII noble metal is selected from platinum and / or palladium. The noble metal catalyst is preferably a Pt-Pd catalyst. Generally, based on the mass of the catalyst, the content of Pt is 0.05wt% - 0.20wt%, and the content of Pd is 0.15wt% - 0.30wt%. The noble metal can use commercial catalysts, such as the FHDA-10 catalyst developed by Fushun Research Institute of Petroleum and Petrochemicals, SINOPEC.

[0020] Furthermore, the gas-phase components discharged from the second reactor are heat-exchanged in a heat exchanger and then enter a high-pressure separator for separation to obtain hydrogenated light components and hydrogen containing hydrogen sulfide.

[0021] Furthermore, the gas-phase components discharged from the second reactor are cooled to 100 - 200°C, preferably 120 - 150°C, in a heat exchanger.

[0022] Furthermore, after the hydrogenated heavy components are mixed with the hydrogenated light components, the separation method used is stripping to obtain a high-smoke-point jet fuel product.

[0023] Furthermore, the hydrogen containing hydrogen sulfide separated by the high-pressure separator can be recycled back to the gas-phase reactor to continue participating in the reaction after removing hydrogen sulfide and ammonia.

[0024] Furthermore, the smoke point of the jet fuel product is ≥25mm.

[0025] Compared with the prior art, the method of the present invention has the following advantages:

[0026] (1) Compared with the conventional fixed-bed jet fuel three-phase hydrogenation technology, the gas-phase reactor in the method of the present invention is under near-atmospheric pressure conditions. Therefore, no pressurization is required during the hydrogen circulation process, eliminating the recycle hydrogen compressor and the involved heat exchange process, significantly reducing energy consumption and construction investment. The entire reaction system is under low-pressure conditions, further reducing energy consumption. From the perspective of the chemical reaction rate, the present invention is a combination of gas-phase and liquid-phase hydrogenation reactions, and the reaction rate is also several orders of magnitude higher than that of gas-liquid-solid three-phase reactions.

[0027] (2) Compared with the two-stage hydrogenation process that uses a highly active catalyst for deep dearomatization, the method of the present invention can eliminate the stripping equipment between the two reactors. Since the material flowing out of the gas-phase reactor is in a high-temperature and low-pressure state, during the slightly pressurized liquefaction process of the compressor, it promotes the synchronous liquefaction and dissolution of hydrogen and oil products. Moreover, due to the different dissolution laws of hydrogen and hydrogen sulfide in oil products, that is, at high temperatures, the solubility of hydrogen is high and the solubility of hydrogen sulfide is low. Therefore, the hydrogen concentration in the liquefied liquid phase is high and the hydrogen sulfide concentration is low, and it can directly enter the subsequent liquid-phase hydrogenation reactor without further stripping. When a noble metal or Ni-based catalyst is loaded in the liquid-phase hydrogenation reactor, it will not have an impact either.

[0028] (3) Through the organic combination of the first reactor and the second reactor, the present invention optimizes the reaction conditions and reaction types of each reactor, and overall reduces the reaction severity and simplifies the process flow. As the deep treatment reactor of the first reactor, the second reactor first removes the substances that are easy to react in the jet fuel raw material under low pressure and high hydrogen-oil ratio conditions. Among the materials flowing out of the gas-phase reactor, the unreacted macromolecules (mainly aromatics) are slightly pressurized and liquefied by the compressor and enter the second reactor as hydrogenation heavy components to undergo deep hydrogenation reactions. Since there is an excess of hydrogen in the first reactor and hydrogen is highly dissolved in the oil product during the pressurization process of the compressor, there is no need to add a hydrogen dissolution device before entering the second hydrogenation reactor. Moreover, since the small-molecule substances have undergone hydrogenation reactions in the first reactor, the pressurization process does not require high pressure, and only a low pressure condition that can liquefy the macromolecules is needed, reducing the reaction severity of the liquid-phase hydrogenation reaction. In addition, since the small-molecule substances have been removed in the first reactor, the aromatic hydrocarbon concentration in the macromolecules entering the second reactor increases significantly, which is conducive to the right shift of the chemical equilibrium and improves the chemical reaction conversion rate. Among the effluents from the first reactor, the reacted small molecules do not liquefy during the pressurization process of the compressor and are liquefied by heat exchange and condensation cooling before entering the high-pressure separator. Such a liquefaction method is conducive to the separation of hydrogen and the raw material, can recover a large amount of hydrogen for recycling, and improves the hydrogen utilization rate. The liquefied small molecules are hydrogenation light components, and after being mixed with the hydrogenation heavy components, they can enter the subsequent separation system. The entire reaction system does not require a hydrogen compressor in the fixed-bed reaction system and a circulating oil pump in the liquid-phase hydrogenation reaction system, reducing the investment cost, simplifying the process flow, improving the reaction efficiency, and reducing the reaction severity. BRIEF DESCRIPTION OF THE DRAWINGS

[0029] Figure 1 It is a process flow diagram for the production of jet fuel by hydrogenation according to the present invention;

[0030] Among them, 1 - jet fuel feedstock and hydrogen, 2 - the first reactor, 3 - the effluent of the first reactor, 4 - compressor, 5 - the second reactor, 6 - hydrotreated heavy components, 7 - gas-phase components, 8 - heat exchanger, 9 - high-pressure separator, 10 - hydrogen containing hydrogen sulfide, 11 - hydrotreated light components, 12 - stripping system, 13 - high smoke point jet fuel. Detailed implementation manners

[0031] The present invention will be further described below in conjunction with embodiments, but it should be understood that the protection scope of the present invention is not limited by the embodiments.

[0032] In the present invention, unless otherwise clearly stated, percentages and percentage contents are by mass.

[0033] In the present invention, the smoke point of jet fuel is measured by the method of GB / T 382, and the analytical instrument is the SH382 jet fuel smoke point tester.

[0034] The following is combined with Figure 1 to describe the process flow of the present invention in detail.

[0035] The jet fuel feedstock and hydrogen 1 enter the first reactor 2, and a gas-phase hydrogenation reaction occurs to obtain the effluent of the first reactor 3, which enters the compressor 4; after being pressurized by the compressor 4, it enters the second reactor 5, where the liquid-phase components continue to undergo a hydrogenation reaction downward to obtain hydrotreated heavy components 6, and the gas-phase components 7 are discharged upward from the second reactor and enter the heat exchanger 8, and then enter the high-pressure separator 9 to be separated into hydrotreated light components 11 and hydrogen containing hydrogen sulfide 10. The hydrotreated heavy components 6 and the hydrotreated light components 11 are mixed and enter the stripping system 12, and finally a high smoke point jet fuel product 13 can be obtained.

[0036] Examples 1 - 3

[0037] Adopt as Figure 1 the process flow schematic diagram. Two 100 mL fixed-bed hydrogenation reactors are connected in series, and a reciprocating compressor is arranged between the reactors. The first reactor is a gas-phase hydrogenation reactor filled with 50 mL of Mo-Ni type diesel hydrogenation catalyst A, and the second reactor is a liquid-phase hydrogenation reactor filled with 50 mL of Pt-Pd series hydrogenation catalyst B. A gas-phase outlet is arranged above the second reactor, which is connected to a heat exchanger (cooled to 130 °C) and a high-pressure separator in sequence. The liquid-phase outlet pipeline at the bottom of the second reactor is connected to the liquid-phase outlet pipeline at the bottom of the high-pressure separator, and they enter the subsequent stripping and fractionation equipment together.

[0038] The properties of the catalysts are shown in Table 1, the properties of the feedstock oil are shown in Table 2, and the reaction process conditions and results are shown in Table 3.

[0039] Comparative Example 1

[0040] Using the conventional single-stage 100 mL fixed-bed hydrotreating process flow, the raw material is the same as that in Example 1. Both the raw material and hydrogen enter the reactor from the top of Reactor 1. 100 mL of Catalyst A is loaded. The reaction process conditions and results are shown in Table 3.

[0041] Comparative Example 2

[0042] Adopt a two-stage hydrogenation process, and set up Hydrogenation Reactor 1 and Hydrogenation Reactor 2. Introduce the same raw material as in Example 1 into Hydrogenation Reactor 1 (loaded with 50 mL of hydrogenation catalyst A) and Hydrogenation Reactor 2 (loaded with 50 mL of hydrogenation catalyst B) in sequence. A stripping device needs to be set between the two hydrogenation reactors.

[0043] Table 1 Physicochemical properties of the catalyst

[0044] Catalyst Number A B Grade FH-40A FHDA-10 Active Metal Ni-Mo Pt-Pd <![CDATA[MoO 3 / Pt, wt%]]> 15 0.10 NiO / Pd, wt% 4.5 0.20 Shape Trifolium Cylinder Diameter, mm 2.0 1.5 <![CDATA[Specific surface area, m 2 ·g -1 > 220 180 <![CDATA[Pore volume, mL·g -1 > 0.26 0.50

[0045] Table 2 Properties of the feedstock oil

[0046]

[0047]

[0048] Table 3 Hydrotreating process conditions and results

[0049] Example 1 Example 2 Example 3 Comparative Example 1 Comparative Example 2 First Reactor (Reactor 1) Pressure, MPa 0.5 0.5 1.0 4.0 4.0 Temperature, °C 240 240 250 250 250 Hydrogen-Oil Volume Ratio, v / v 500 400 500 500 500 <![CDATA[Space velocity, h -1 > 3.0 3.0 3.0 1.5 3.0 Second Reactor (Reactor 2) Pressure, MPa 3.0 4.0 4.0 — 4.0 Temperature, °C 120 130 140 — 140 <![CDATA[Space velocity, h -1 > 1.0 1.0 1.0 — 3.0 Properties of High Smoke Point Jet Fuel Product Sulfur, μg / g 121 127 86 133 99 Smoke Point, mm 25.6 25.9 27.6 23.3 24.9

[0050] As can be seen from Table 3, for the conventional fixed-bed hydrogenation technology using heavy cut jet fuel as the raw material, when the operating conditions of an industrial plant are adopted, due to insufficient aromatic saturation, the smoke point index cannot meet the requirements. When using the two-stage hydrogenation technology, hydrogen sulfide needs to be stripped off after the first-stage hydrogenation, and then the whole fraction enters Hydrogenation Reactor 2. Under the same processing capacity and catalyst loading, it is equivalent to increasing the volumetric space velocity of the aromatic hydrogenation part, which affects the reaction effect. The present invention can achieve improving the hydrogenation reaction efficiency under low reaction severity and a simplified process flow, and produce high-quality jet fuel products using heavy cut jet fuel as the raw material.

Claims

1. A method for producing jet fuel by hydrogenation, characterized in that, the method comprises the following steps: (1) The jet fuel raw material and hydrogen enter the first reactor for gas-phase hydrogenation reaction to obtain the effluent of the first reactor; wherein, under the conditions of low pressure and large hydrogen-oil ratio, the substances in the jet fuel raw material that are easy to react are removed first. In the material flowing out of the first reactor, the unreacted macromolecules are pressurized and liquefied by a compressor and enter the second reactor for deep hydrogenation reaction; (2) The effluent of the first reactor enters the second reactor after being pressurized. Among them, the gas-phase components are discharged upward from the second reactor, and the liquid-phase components undergo liquid-phase hydrodearomatization reaction downward to obtain hydrogenated heavy components; (3) The gas-phase components discharged from the second reactor are separated to obtain hydrogenated light components and hydrogen containing hydrogen sulfide; (4) After the hydrogenated heavy components and the hydrogenated light components are mixed, they are separated to obtain jet fuel products, and the smoke point of the jet fuel products is ≥25 mm; The operating conditions of the first reactor are as follows: the pressure is 0.1 - 2.0 MPa, the hydrogen-oil volume ratio is 100 - 1000, the temperature is 150 - 300 °C, and the volume space velocity is 0.1 - 6.0 h -1 ; The operating conditions of the second reactor are as follows: the pressure is 1.5 - 6.0 MPa, the temperature is 100 - 230 °C, and the volume space velocity is 0.5 - 4.0 h -1 .

2. The method according to claim 1, characterized in that, the properties of the jet fuel raw material are as follows: the initial boiling point is 100°C to 150°C, the final boiling point is 280°C to 320°C, the smoke point is ≥10 mm, the S content is 600 to 3000 μg / g, and the N content is 5 to 300 μg / g.

3. The method according to claim 2, characterized in that, the smoke point of the jet fuel raw material is 15 to 22 mm.

4. The method according to claim 1, characterized in that, both the first reactor and the second reactor are fixed-bed reactors; the second reactor uses a fixed-bed reactor provided with a flash zone.

5. The method according to claim 1, characterized in that, The operating conditions of the first reactor are as follows: the pressure is 0.5 - 2.0 MPa; the hydrogen-oil volume ratio is 200 - 600; the temperature is 200 - 280 °C; the volumetric space velocity is 1.0 - 4.0 h -1 .

6. The method according to claim 1, characterized in that, The operating conditions of the second reactor are as follows: the pressure is 1.5 to 5.0 MPa; the temperature is 100 to 160 °C; the volume space velocity is 1.0 to 3.0 h -1 .

7. The method according to claim 1, characterized in that, the pressure of the second reactor is at least 0.5 MPa higher than the pressure of the first reactor.

8. The method according to claim 6, characterized in that, the pressure of the second reactor is 1.0 to 6.0 MPa higher than the pressure of the first reactor.

9. The method according to claim 6, characterized in that, the pressure of the second reactor is 2.0 to 4.0 MPa higher than the pressure of the first reactor.

10. The method according to claim 1, characterized in that, the gas-phase components discharged from the second reactor are heat-exchanged through a heat exchanger and then enter a high-pressure separator for separation to obtain hydrogenated light components and hydrogen containing hydrogen sulfide.

11. The method according to claim 10, characterized in that, the gas-phase components discharged from the second reactor are cooled to 100 to 200°C through a heat exchanger.

12. The method according to claim 11, characterized in that, the gas-phase components discharged from the second reactor are cooled to 120 to 150°C through a heat exchanger.

13. The method according to claim 1, characterized in that, the catalyst loaded in the first reactor includes a hydrofining catalyst, and the catalyst loaded in the second reactor is a hydrodearomatization catalyst.

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