A method for producing a special oil

Abstract

The application discloses a production method of special oil. The method comprises the following steps: (1) wax oil raw material and hydrogen enter a first reactor to perform a gas phase desulfurization reaction; (2) first reactor effluent enters a second reactor after being pressurized, wherein gas phase components are discharged upwards from the second reactor to remove hydrogen sulfide and obtain a hydrogenated light component, and liquid phase components perform a liquid phase isomerization reaction downwards to obtain a hydrogenated heavy component which is discharged from the bottom of the second reactor; and (3) the hydrogenated heavy component is mixed with the hydrogenated light component to enter a third reactor to perform a deep de-aromatics reaction, and finally special oil is obtained. The method can achieve the effects of deep desulfurization and de-aromatics, and obtain special oil with excellent comprehensive properties; meanwhile, the method simplifies the process flow as a whole, reduces reaction severity, and improves chemical reaction efficiency.

Description

TECHNICAL FIELD

[0001] The present application belongs to the field of oil refining and chemical industry, and particularly relates to a production method of special oil products. BACKGROUND

[0002] With the advent of the new energy era and the structural adjustment of oil refining enterprises, the demand for fuel oil increases at a slower rate, and the demand for high value-added products such as special oil products and chemical products is increasing. "Oil to special" and "oil conversion" have become an important direction for the technical transformation and development of the oil refining industry, which not only maximizes the utilization of crude oil, but also solves the defects of single oil refining products and low added value. For example, white oil and lubricating oil base oil are produced from wax oil. In the hydrogenation reaction process, desulfurization and de-aromatic reaction mainly occur, and the sulfides need to be removed by hydrofining first, and the saturated aromatic hydrocarbons need to be removed, and then isomerization dewaxing and make-up refining are carried out to reduce the condensation point of the raw material and deeply remove the aromatic hydrocarbons to improve the stability of the product. The refined product can be obtained after fractionation, such as lubricating oil base oil or industrial white oil, and finally noble metal deep refining can be carried out to obtain food / medical grade white oil. The hydrofining process uses a fixed bed reactor and is carried out at a relatively high temperature and pressure. In addition, since the isomerization dewaxing reactor is generally filled with noble metal catalyst, the material after hydrofining needs to be first stripped of hydrogen sulfide in a stripping tower and then introduced into the isomerization reactor, so the energy consumption for separation is also high.

[0003] In addition, from the reaction mechanism of deep desulfurization and deep de-aromatics, there is a big difference in the requirements for the reaction environment. For desulfurization reaction, the removal of small molecule sulfur mainly follows the direct desulfurization route, that is, hydrogenolysis desulfurization. The large molecular sulfides with low reactivity follow the hydrogenation desulfurization reaction route, that is, first hydrogenation of aromatic ring, and then hydrogenolysis desulfurization. Therefore, in the upper part of the reactor, the relatively low temperature and high hydrogen partial pressure environment, the hydrogenolysis of small molecule sulfides (endothermic) and the hydrogenation of aromatic ring-containing substances (large exothermic) mainly occur. In the lower part of the reactor, the accumulation of heat and hydrogen sulfide makes the reaction environment high temperature and low hydrogen partial pressure, which is beneficial to the further hydrogenolysis of large molecular sulfides after hydrogenation, but is very unfavorable for the further saturation of aromatic hydrocarbons, because it is severely limited by the thermodynamics of aromatic hydrocarbon hydrogenation. Especially at the end of the reaction, the catalyst activity decays, and there is no better operation method except increasing the temperature, which will further affect the de-aromatic effect. In addition to the different requirements for the reaction environment, the competitive adsorption of aromatic hydrocarbons on the catalyst surface also has an inhibitory effect on deep desulfurization, making it difficult for traditional hydrogenation technology to meet the dual requirements of ultra-deep desulfurization and efficient saturation of aromatic hydrocarbons.

[0004] For the prior art, on the basis of deep desulfurization, one is to adopt the existing catalyst system, reduce the reaction space velocity, that is, reduce the processing capacity or increase the method of the reactor, but it is not very reasonable from the economic point of view. Two is to adopt two-stage process technology, after conventional hydrogenation, the generated oil is stripped to remove hydrogen sulfide, and then enters the noble metal hydrogenation reactor, which will greatly increase the catalyst use cost and process complexity, which is still not the best solution.

[0005] CN112410068A discloses a wax oil hydrocracking device and a method for producing 5# industrial white oil, which comprises a delivery pump, a hydrofining reactor and a hydrocracking reactor. The cracking reaction products are subjected to gas-liquid separation, stripping of hydrogen sulfide, flash distillation and fractionation in a cold and hot high-low separator to obtain naphtha, aviation coal, industrial white oil and tail oil. However, since the reaction system relies on hydrocracking for de-aromatization, the reaction pressure is high, resulting in high energy consumption.

[0006] CN111154506A discloses a method for producing white oil by hydrogenation of coal tar, which comprises mixing coal tar with generated oil of a hydrocracking unit, then sequentially subjecting to acid washing treatment and water washing treatment to obtain feed of a hydrofining unit, and then performing hydrogenation reaction. The hydrogenation reaction products are subjected to fractionation to obtain naphtha, 3# crude white oil, 5# crude white oil and hydrogenation tail oil. The method needs to combine a hydrocracking unit and a deep hydrofining unit, which has high pressure and high energy consumption. SUMMARY

[0007] In view of the deficiencies of the prior art, the present application provides a production method of special oil. The method of the present application can achieve the effects of deep desulfurization and de-aromatics, and obtain special oil with excellent comprehensive properties. At the same time, the process flow is simplified as a whole, the reaction severity is reduced, and the chemical reaction efficiency is improved.

[0008] The present application provides a production method of special oil, which comprises the following steps:

[0009] (1) The wax oil raw material and hydrogen gas enter the first reactor to perform gas phase desulfurization reaction;

[0010] (2) The first reactor effluent is pressurized and then enters the second reactor, wherein the gas phase components are discharged upward from the second reactor to remove hydrogen sulfide, and the liquid phase components are subjected to liquid phase isomerization reaction downward to obtain hydrogenation heavy components which are discharged from the bottom of the second reactor;

[0011] (3) The hydrogenation heavy components are mixed with the hydrogenation light components and then enter the third reactor to perform deep de-aromatics reaction, and finally obtain special oil.

[0012] Further, the properties of the wax oil feedstock are as follows: the initial boiling point is 150-200°C, the final boiling point is 450-550°C; the S content is ≯ 15000 μg / g, preferably 3000-7000 μg / g, the N content is ≯ 1000 μg / g, preferably 200-800 μg / g, the aromatic hydrocarbon content is ≯ 70 wt%, preferably 15 wt%-40 wt%. For example, the S content of the wax oil feedstock can be exemplified but is not limited to 3000 μg / g, 5000 μg / g, 6500 μg / g, etc., the N content can be exemplified but is not limited to 200 μg / g, 400 μg / g, 700 μg / g, etc., and the aromatic hydrocarbon content can be exemplified but is not limited to 15 wt%, 20 wt%, 25 wt%, etc.

[0013] Further, the wax oil feedstock can be one or more of atmospheric straight-run wax oil, vacuum straight-run wax oil, coking wax oil, etc.

[0014] Further, the first reactor, the second reactor and the third reactor are all fixed bed reactors. The second reactor preferably uses a fixed bed reactor provided with a flash zone.

[0015] Further, the second reactor is provided with a flash zone, no catalyst is loaded in the flash zone and above, and the reaction zone is below the flash zone. The first reactor effluent is fed to the flash zone of the second reactor after being pressurized, the obtained gas phase component is discharged upward from the second reactor, and the obtained liquid phase component is subjected to isomerization and condensation reduction reaction downward, and the obtained hydrogenated heavy component is discharged from the bottom of the second reactor.

[0016] Further, the catalyst loaded in the first reactor can use a conventional catalyst with hydrodesulfurization function, preferably a Mo-Co type hydrodesulfurization catalyst. Generally, an alumina-based carrier is used, and the content of Mo calculated as molybdenum oxide is 15 wt%-30 wt% and the content of Co calculated as cobalt oxide is 2 wt%-6 wt% based on the mass of the catalyst. The catalyst can also contain an auxiliary component, such as at least one of phosphorus, silicon, boron, magnesium, fluorine, etc., and the mass content in the catalyst is generally 6 wt% or less. For example, the Mo-Co type diesel deep desulfurization catalyst developed by Fushun Research Institute of Petroleum and Petrochemicals (FRIPP). For example, FHUDS-5, FHUDS-7, etc.

[0017] Further, the catalyst loaded in the second reactor can use a catalyst with isomerization and condensation reduction function, such as a Ni type catalyst. Generally, an alumina added molecular sieve is used as the carrier, and the content of Ni is 0.2 wt%-2.0 wt% based on the mass of the catalyst. The Ni type catalyst is, for example, the FDW-3 developed by Fushun Research Institute of Petroleum and Petrochemicals (FRIPP) and the RIDW developed by Research Institute of Petroleum and Petrochemicals.

[0018] Further, the catalyst filled in the third reactor can adopt a catalyst with deep hydrogenation function, such as a noble metal catalyst or a non-noble metal catalyst. The noble metal catalyst is a Pt-Pd type catalyst, such as the FHDA-10 catalyst developed by FRIPP. The non-noble metal catalyst is a Mo-Ni type catalyst, which generally adopts an alumina-based carrier, and the content of Mo, calculated as molybdenum oxide, is 15wt% to 30wt%, and the content of Ni, calculated as cobalt oxide, is 2wt% to 5wt%, based on the mass of the catalyst. The catalyst can also contain an auxiliary component, such as at least one of phosphorus, silicon, boron, magnesium, fluorine, etc., and the mass content in the catalyst is generally 6wt% or less. The Mo-Ni type catalyst is, for example, FHUDS-10, FHUDS-6, FHUDS-8, etc. developed by FRIPP.

[0019] Further, 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 temperature is 300 to 400°C, preferably 320 to 370°C; the hydrogen / oil volume ratio is 100 to 900, preferably 400 to 700; and the volume space velocity is 0.5 to 3.0 h -1 , preferably 0.8 to 2.0 h -1 .

[0020] Further, the first reactor effluent is pressurized by a compressor, which can be a conventional commercial compressor, such as a reciprocating compressor or a centrifugal compressor. The pressurization can ensure normal feeding of the second reactor and meet the operating pressure requirements of the second reactor.

[0021] Further, the operating conditions of the second reactor are as follows: the pressure is 4.0 to 15.0 MPa, preferably 6.0 to 10.0 MPa; the temperature is 250 to 450°C, preferably 320 to 380°C; and the volume space velocity is 0.1 to 3.0 h -1 , preferably 0.5 to 1.5 h -1 .

[0022] Further, the operating conditions of the third reactor are as follows: the pressure is 4.0 to 10.0 MPa, preferably 4.0 to 8.0 MPa; the temperature is 260 to 400°C, preferably 300 to 360°C; and the volume space velocity is 0.1 to 4.0 h -1 , preferably 1.0 to 3.0 h -1 .

[0023] Further, the pressure of the second reactor is at least 2.0 MPa higher than that of the first reactor, preferably 4.0 to 8.0 MPa higher.

[0024] Further, the gas phase component discharged from the second reactor is cooled by a heat exchanger and then enters a high-pressure separator to separate a hydrogen gas containing hydrogen sulfide and a hydrogenated light component.

[0025] Further, the gas phase component discharged from the second reactor is cooled by a heat exchanger to 100-200°C, preferably 120-150°C.

[0026] Further, the hydrogenated heavy component is mixed with the hydrogenated light component and then enters a third reactor to undergo a deep de-aromatic reaction, and a special oil product with low sulfur and low aromatic content is obtained at the bottom of the reactor.

[0027] Further, the special oil product has an aromatic content of 5wt% or less, preferably 2wt%-5wt%, and an S content of 10μg / g or less, preferably 0.1-5μg / g.

[0028] Further, the special oil product can be further fractionated or processed according to actual needs to obtain different types of special oil products.

[0029] Compared with the prior art, the method has the following advantages:

[0030] (1) The method controls the desulfurization, isomerization and de-aromatic reactions to occur in three reactors respectively, avoiding the defects that the reactions are placed in the same reaction system and the reaction conditions are difficult to be compatible. The inventors have found that by controlling the reaction conditions of the first reactor, the aromatic ring-containing substances are not easy to be adsorbed on the surface of the catalyst, the competitive adsorption of aromatic substances is greatly reduced, and the targeted desulfurization reaction is facilitated. The second reactor is a liquid phase reaction, and the mass transfer efficiency is higher, so that better reaction effect can be obtained. In addition, since the macromolecular sulfides can be deeply removed in the first reactor, only a deep de-aromatic reaction needs to occur in the third reactor, avoiding the limitation of different requirements for reaction environment and catalyst type due to the removal of macromolecular sulfur and de-aromatic reaction.

[0031] (2) In order to avoid the toxic effect of hydrogen sulfide in the hydrogenation desulfurization stream on the noble metal isomerization catalyst in the subsequent process, the hydrogen sulfide needs to be stripped by a stripping tower in the conventional process. In the present application, the outflow of the first reactor is pressurized under high temperature conditions to promote the synchronous liquefaction of hydrogen and oil. Since the solubility of hydrogen is high and the solubility of hydrogen sulfide is low under high temperature conditions, the concentration of hydrogen is high and the concentration of hydrogen sulfide is low in the liquefied liquid phase. When entering the subsequent second reactor, the influence of hydrogen sulfide on the activity of the catalyst is avoided, and the setting of stripping hydrogen sulfide between the two reactors is also saved.

[0032] (3) Since the first reactor is operated at low pressure and the second and third reactors are operated at medium-high pressure, the hydrogen separated from the high fraction can directly enter the device during the recycling process without the need to set up a hydrogen compressor, thereby greatly reducing the investment of the device.

[0033] (4) The present application realizes the production of special oil products from waxy oil under more moderate operating conditions. The reaction pressure of the gas phase reactor is lower than that of conventional hydrogenation technology, thereby greatly reducing the energy consumption of the device. In the first reactor effluent, the reacted small molecules are not liquefied during the pressurization process of the compressor, but are liquefied through the heat exchanger and the high-pressure separator. Such liquefaction means is beneficial to the separation of hydrogen and raw materials, and a large amount of hydrogen can be recycled and used, thereby improving the hydrogen utilization rate. The hydrogenated light components separated by the high-pressure separator can be mixed with the hydrogenated heavy components and then enter the subsequent deep refining system. BRIEF DESCRIPTION OF DRAWINGS

[0034] Figure 1 Flowchart for the production method of special oil products in Examples 1-3;

[0035] Among them, waxy oil raw material and hydrogen-1; first reactor-2, first reactor effluent-3; compressor-4; second reactor-5; hydrogenated heavy component-6; gas phase component-7; heat exchanger-8; high-pressure separator-9; hydrogenated light component-10; hydrogen containing hydrogen sulfide-11; third reactor-12; special oil product-13. DETAILED DESCRIPTION

[0036] The present application will be further described in conjunction with the examples. It should be understood that the protection scope of the present application is not limited by the examples.

[0037] In the present application, unless otherwise explicitly stated, the percentage and percentage content are measured by mass.

[0038] The following will be described in conjunction with Figure 1 The process flow of the present application will be described in detail.

[0039] The waxy oil raw material and hydrogen 1 enter the first reactor 2 to undergo gas phase desulfurization reaction to obtain the first reactor effluent 3; enter the compressor 4, and after being pressurized by the compressor 4, enter the second reactor 5, wherein the liquid phase component enters the reaction zone downward to undergo isomerization and condensation reaction to obtain the hydrogenated heavy component 6, and the gas phase component 7 is discharged upward from the second reactor, enters the heat exchanger 8, and then enters the high-pressure separator 9 to be separated into the hydrogenated light component 10 and the hydrogen containing hydrogen sulfide 11. The hydrogenated heavy component 6 and the hydrogenated light component 10 are mixed to enter the third reactor 12, and finally the special oil product 13 is obtained.

[0040] Examples 1-3

[0041] As described above, the present application adopts the method of Figure 1The flow diagram of the process is shown in Figure 1. Three 100 mL fixed bed hydrogenation reactors are connected in series, which are the first reactor, the second reactor and the third reactor. A conventional power reciprocating compressor is arranged between the first reactor and the second reactor. The first reactor is a gas phase hydrogenation reactor, which is filled with 50 mL of Mo-Co type diesel hydrogenation catalyst A. The second reactor is a liquid phase hydrogenation reactor, which is filled with 50 mL of Ni type isomerization and pour point depressing catalyst B. A gas phase outlet is arranged above the second reactor, which is connected with 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 with the liquid phase outlet pipeline at the bottom of the high pressure separator, which enters the third reactor together. The third reactor is filled with 50 mL of Mo-Ni type deep hydrogenation catalyst C, and finally the special oil product is obtained.

[0042] The atmospheric straight-run wax oil is used as the raw material. The catalyst properties are shown in Table 1, the raw oil properties are shown in Table 2, and the reaction process conditions and results are shown in Table 3.

[0043] Comparative Example 1

[0044] The process flow of the combined hydrogenation of conventional wax oil to produce special oil products is used, and three fixed bed hydrogenation reactors are connected in series, which are reactor 1, reactor 2 and reactor 3. A stripping tower is arranged between reactor 1 and reactor 2 for stripping hydrogen sulfide to avoid its influence on the activity of the noble metal catalyst. The three reactors do not control the phase state of the fluid in the reactor through other equipment or operating conditions, so the gas-liquid-solid three-phase reaction occurs in the reactor. Reactor 1 is filled with 50 mL of Mo-Co type diesel hydrogenation catalyst A, reactor 2 is filled with 50 mL of Ni type isomerization and pour point depressing catalyst B, and reactor 3 is filled with 50 mL of Mo-Ni type deep hydrogenation catalyst C, and finally the special oil product is obtained.

[0045] The raw material and catalyst properties are the same as those in Examples 1-3, and the reaction process conditions and results are shown in Table 3.

[0046] Comparative Example 2

[0047] The process flow of the combined hydrogenation of wax oil to produce special oil products is used, and three fixed bed hydrogenation reactors are connected in series, which are reactor 1, reactor 2 and reactor 3. The second reactor is a liquid phase hydrogenation reactor, and hydrogen is dissolved in the oil product through a hydrogen mixer before entering the second reactor to enter the reactor 2 in a liquid phase state, and a liquid-solid two-phase reaction occurs. A stripping tower is arranged between reactor 1 and reactor 2, reactor 1 is filled with 50 mL of Mo-Co type diesel hydrogenation catalyst A, reactor 2 is filled with 50 mL of Ni type isomerization and pour point depressing catalyst B, and reactor 3 is filled with 50 mL of Mo-Ni type deep hydrogenation catalyst C, and finally the special oil product is obtained.

[0048] The raw material and catalyst properties are the same as those in Examples 1-3, and the reaction process conditions and results are shown in Table 3.

[0049] Table 1 Catalyst physicochemical properties

[0050]

[0051]

[0052] Table 2 Properties of raw oil

[0053] Oil properties Straight-run gas oil Density (20°C), g-cm -3 ]] 0.875 Distillation range, °C 180～513 [S, pg g -1 ]] 3900 [N, pg g -1 ]] 461 Aromatics, wt% 35 Pour point, °C +33 Kinematic viscosity (100°C), mm 2 ·s -1 ]]> 4.757

[0054] Table 3 Hydrogenation process conditions and results

[0055]

[0056]

[0057] As can be seen from Table 3, the comparative example 1 uses the conventional wax oil hydrogenation fixed bed combination technology, the first hydrogenation refining reactor, i.e. reactor 1 needs to use a higher pressure, in order to meet the requirements of the noble metal isomerization and condensation reactor feed, a stripping tower needs to be set up to strip hydrogen sulfide, and the process flow is relatively complex. In addition, due to the simultaneous occurrence of desulfurization and de-aromatic reaction in reactor 1, the competitive adsorption of aromatic hydrocarbons will affect the desulfurization effect, leading to the further removal of sulfides which can only be completed in the deep refining process of reactor 3, thereby affecting the effect of deep de-aromatic.

[0058] In order to obtain better isomerization reaction effect, the comparative example 2 uses a liquid phase hydrogenation reactor in the conventional wax oil hydrogenation fixed bed combination technology in reactor 2, but due to the liquid phase of the reactor 1 effluent, the hydrogen gas cannot be dissolved by using the phase change process of gas, so a hydrogen gas mixer needs to be used to dissolve hydrogen, leading to an increase in the investment of the device.

Claims

1. A method for producing a specialty oil product, characterized by, The method comprises the following steps: (1) wax oil raw material and hydrogen enter a first reactor to perform a gas phase desulfurization reaction; (2) the first reactor effluent enters a second reactor after being pressurized, wherein a gas phase component is discharged upward from the second reactor to remove hydrogen sulfide and obtain a hydrogenation light component, and a liquid phase component performs a liquid phase isomerization reaction downward to obtain a hydrogenation heavy component which is discharged from the bottom of the second reactor; (3) the hydrogenation heavy component is mixed with the hydrogenation light component and enters a third reactor to perform a deep de-aromatic reaction to obtain a special oil product; The first reactor, the second reactor and the third reactor are all fixed bed reactors, and the second reactor is a fixed bed reactor provided with a flash zone; The operating conditions of the first reactor are as follows: pressure 0.1-3.0 MPa, temperature 300-400℃, hydrogen / oil volume ratio 100-900, volume space velocity 0.5-3.0 h -1 ; The operating conditions of the second reactor are as follows: pressure is 4.0-15.0 MPa, temperature is 200-400 ℃, volume space velocity is 0.1-3.0 h -1 ; The pressure of the second reactor is at least 2.0 MPa higher than that of the first reactor; The operating conditions of the third reactor are as follows: pressure 4.0-10.0 MPa, temperature 260-400℃, volume space velocity 0.1-4.0 h -1 .

2. The method of claim 1, wherein, The wax oil raw material has the following properties: an initial boiling point of 150-200 DEG C, a final boiling point of 450-550 DEG C, an S content of ≤15000 μg / g, an N content of ≤1000 μg / g and an aromatic hydrocarbon content of ≤70 wt%.

3. The method of claim 2, wherein, The wax oil raw material has the following properties: an S content of 3000-7000 μg / g, an N content of 200-800 μg / g and an aromatic hydrocarbon content of 15 wt%-40 wt%.

4. The method of claim 1, wherein, The operating conditions of the first reactor are as follows: pressure 0.5-2.0 MPa, temperature 320-370℃, hydrogen / oil volume ratio 400-700, volume space velocity 0.8-2.0 h -1 .

5. The method of claim 1, wherein, The operating conditions of the second reactor are as follows: pressure is 6.0-10.0 MPa, temperature is 280-340℃, volume space velocity is 0.5-1.5h -1 .

6. The method according to claim 1 or 4 or 5, characterized in that, The pressure of the second reactor is at least 4.0-8.0 MPa higher than that of the first reactor.

7. The method of claim 1, wherein, The operating conditions of the third reactor are as follows: the pressure is 4.0-8.0 MPa, the temperature is 300-360 ℃, the volume space velocity is 1.0-3.0 h -1 .

8. The method of claim 1, wherein, The gas phase component discharged from the second reactor is heated by a heat exchanger and then enters a high-pressure separator to be separated to obtain a hydrogenation light component and hydrogen containing hydrogen sulfide.

9. The method of claim 1, wherein, The hydrogenation heavy component is mixed with the hydrogenation light component and then enters the third reactor to perform a deep de-aromatic reaction, and a special oil product is obtained at the bottom of the reactor.

10. The method according to claim 1 or 9, characterized in that, The aromatic hydrocarbon content of the special oil product is ≤5 wt%, and the S content is ≤10 μg / g.

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