CN109777487B - Refinery gas combined processing method - Google Patents
Refinery gas combined processing method Download PDFInfo
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- CN109777487B CN109777487B CN201711118776.3A CN201711118776A CN109777487B CN 109777487 B CN109777487 B CN 109777487B CN 201711118776 A CN201711118776 A CN 201711118776A CN 109777487 B CN109777487 B CN 109777487B
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Abstract
The invention discloses a refinery gas combined processing method, which comprises the following steps: (a) mixing wax oil raw oil and circulating oil with hydrogen in hydrogen dissolving equipment, and then entering a hydrotreating catalyst bed layer in a hydrotreating reactor to react under the condition of liquid-phase hydrogenation operation; (b) mixing the reactant flow obtained in the step (a) with refinery gas and hydrogen in a gas dissolving device, and then allowing the mixture to enter a hydrogenation catalyst bed layer in a supplementary hydrogenation reactor to react under the liquid-phase hydrogenation operation condition; (c) separating the hydrogenation reaction effluent in the step (b) into a gas phase and a liquid phase, and fractionating the liquid phase to obtain naphtha, diesel oil and hydrotreated heavy distillate oil; (d) mixing part or all of the diesel oil and/or the hydrotreating heavy fraction obtained in the step (c) with hydrogen, allowing the mixture to enter a hydrocracking catalyst bed layer in a hydrocracking reactor for reaction, and fractionating the separated liquid in a fractionating tower to obtain naphtha, aviation kerosene, diesel oil and tail oil. The method can simultaneously hydrotreat refinery gas and high-quality hydrocracking products.
Description
Technical Field
The invention belongs to a hydrogenation process of an oil refining technology, relates to a refinery gas combined processing method, and particularly relates to a hydrogenation combined method for hydrotreating refinery gas and producing high-quality hydrocracking products.
Background
With the stricter environmental regulations, the quality requirement of motor fuel is higher and higher. The most outstanding characteristic of hydrocracking technology is that clean motor fuels such as clean diesel oil, high-quality aviation kerosene and the like without sulfur, low aromatic hydrocarbon and high cetane number, and high-quality petrochemical raw materials such as light naphtha, heavy naphtha, tail oil and the like can be directly produced from inferior and heavy raw materials. But also has the characteristics of high production flexibility, high liquid product yield and the like. With the upsizing of the hydrocracking device, the hydrocracking technology will be further developed and applied. From the viewpoint of processing flow, two-stage hydrocracking process was first developed and widely used. The two-stage hydrocracking technology mainly comprises a first-stage hydrocracking pretreatment reaction zone and a second-stage hydrocracking reaction zone, wherein in the first-stage hydrocracking pretreatment reaction zone, raw oil and hydrogen undergo reactions such as desulfurization, denitrification, deoxidation and olefin aromatic hydrocarbon hydrogenation saturation, the first-stage generated oil undergoes oil-gas separation, and liquid enters the second-stage hydrocracking reaction zone after being subjected to steam stripping and/or fractionation to undergo hydrocracking and olefin aromatic hydrocarbon hydrogenation saturation reactions, and a small amount of reactions such as hydrodesulfurization, denitrification, deoxidation and the like are continuously performed.
The liquid-phase wax oil hydrocracking technology can meet the requirement of clean diesel oil production under the condition of greatly reducing energy consumption. US6213835, US6428686 and CN103797093B disclose a hydrogenation process of pre-dissolved hydrogen, which all dissolve hydrogen in wax oil raw material to carry out hydrogenation reaction, and do not utilize the hydrogen left in the reaction, and separate and treat it separately.
Refinery gases generally include dry gases, liquefied gases, and the like, and have various paths for their use. The main application comprises that dry gas is hydrogenated and then used as a raw material for preparing ethylene by steam cracking, liquefied gas is hydrogenated and then used as a raw material for preparing ethylene by steam cracking, a raw material for synthesizing maleic anhydride, liquefied gas for vehicles and the like. In the existing refinery gas hydrogenation technology, CN201410271572.3 discloses a coking dry gas hydrogenation catalyst and a catalyst grading method. The method only solves the problem of controlling the reaction temperature during the hydrogenation of the coking dry gas, but the temperature rise in the reaction process is large. CN201010221244.4 discloses a method for preparing ethylene cracking material by hydrogenation of liquefied petroleum gas, which comprises two reactors, a cooling facility is arranged between the reactors, and CN201310628425.2 discloses a high-temperature hydrogenation purification process of liquefied petroleum gas, wherein olefin saturation and hydrogenation are performed by hydrogenation to remove impurities. As is well known, the hydrogenation reaction of unsaturated hydrocarbons such as olefin, diene, alkyne and the like is a strong exothermic reaction, the temperature rise in the gas hydrogenation process is very large, generally 100-200 ℃, the balance of the hydrogenation reaction is damaged along with the temperature rise, and the generation of carbon deposition is seriously increased, so that the service cycle of the catalyst is reduced.
CN201010221263.7 discloses a liquefied petroleum gas-coker gasoline hydrogenation combination process method, which is a combination method, but not a liquid phase hydrogenation method, the coker gasoline is firstly mixed with hydrogen to carry out fixed bed hydrogenation reaction, and a hydrogenation product and liquefied gas are mixed and enter another reactor, so that the problem of hydrogenation temperature rise of the liquefied gas is only solved.
In summary, in the prior art, the hydrotreating process of refinery gas is a gas phase reaction, the wax oil hydrogenation is a liquid phase reaction, and the reaction types of the two are completely different, so the combined method of the refinery gas hydrotreating and the wax oil liquid phase hydrocracking is rarely reported.
Disclosure of Invention
Aiming at the defects of the prior art, the invention provides a hydrogenation combination processing method. The method can simultaneously hydrotreat refinery gas and produce high-quality hydrocracking products. The utilization efficiency of hydrogen is improved on the premise of ensuring the quality of the hydrocracking products, the problem of temperature rise in the hydrotreating process of refinery gas is effectively solved, the equipment investment is reduced overall, and the operation energy consumption is reduced.
The invention relates to a refinery gas combined processing method, which comprises the following steps:
(a) mixing the wax oil raw oil and the circulating oil with hydrogen in hydrogen dissolving equipment, and then entering a hydrotreating catalyst bed layer in a wax oil hydrotreating reactor to react under the liquid-phase hydrogenation operation condition;
(b) mixing the reactant flow obtained in the step (a) with refinery gas and hydrogen in a gas dissolving device, and then allowing the mixture to enter a hydrogenation catalyst bed layer in a supplementary hydrogenation reactor to react under the liquid-phase hydrogenation operation condition;
(c) separating the hydrogenation reaction effluent in the step (b) into a gas phase and a liquid phase, continuously separating the gas phase obtained by separation after removing hydrogen sulfide to obtain hydrogen and hydrotreated refinery gas, fractionating the liquid phase obtained by separation to obtain naphtha, diesel oil and hydrotreated heavy distillate oil, and returning part of the hydrogenation reaction effluent obtained in the step (a) and/or part of the hydrogenation reaction material flow obtained in the step (b) and/or part of the liquid phase obtained by separation of a high-pressure separator as circulating oil to hydrogen dissolving equipment;
(d) mixing part or all of the diesel oil and/or the hydrotreated heavy fraction obtained in the step (c) with hydrogen, allowing the mixture to enter a hydrocracking catalyst bed layer in a hydrocracking reactor for reaction, separating reactant streams in a high-pressure separator, recycling the separated gas, and fractionating the separated liquid in a fractionating tower to obtain naphtha, aviation kerosene, diesel oil and tail oil.
In the above method, the wax oil raw oil used may include VGO, CGO, HGO, HLCO, DAO, etc. obtained from petroleum, coal tar, coal liquefied oil, etc., and may be a raw oil, a mixed raw oil composed of several raw materials, or light distillate oil blended with LCO, etc.
In the method, the hydrotreating operation condition is generally that the reaction pressure is 3.0-20.0 MPa, and the volume space velocity of the wax oil raw material oil is 0.2h-1~8.0h-1The average reaction temperature is 180-450 ℃, and the ratio of the circulating oil to the wax oil raw oil is 0.5: 1-10: 1; the preferable operation conditions are that the reaction pressure is 4.0 MPa-18.0 MPa, and the volume space velocity of the wax oil raw material oil is 0.5h-1~6.0h-1The average reaction temperature is 200-440 ℃, and the ratio of the circulating oil to the wax oil raw oil is 0.6: 1-8: 1.
In the method, the supplementary hydrotreating operation condition is generally that the reaction pressure is 3.0-20.0 MPa, and the volume space velocity of the wax oil raw material oil is 0.5h-1~40.0h-1The average reaction temperature is 180-450 ℃; the preferable operation conditions are that the reaction pressure is 4.0 MPa-18.0 MPa, and the volume space velocity of the wax oil raw material oil is 0.8h-1~30.0h-1The average reaction temperature is 200-440 ℃.
In the method, the hydrogenation active component in the hydrogenation catalyst is one or more of Co, Mo, W and Ni, the weight content of the hydrogenation active component is 5-70% in terms of oxide, the carrier of the hydrogenation catalyst is generally alumina, amorphous silicon aluminum, silicon oxide, titanium oxide and the like, and other auxiliary agents such as P, Si, B, Ti, Zr and the like can be simultaneously contained. The catalyst may be used commercially or may be prepared by methods known in the art. The hydrogenation active component is a catalyst in an oxidation state, and is subjected to conventional vulcanization treatment before use, so that the hydrogenation active component is converted into a vulcanization state. The commercial hydrogenation catalysts mainly comprise hydrogenation catalysts such as 3926, 3936, CH-20, FF-14, FF-18, FF-24, FF-26, FF-36, FF-46, FF-56, FH-98, FH-UDS and FZC-41 developed by the Fushu petrochemical research institute (FRIPP), hydrogenation catalysts such as HR-406, HR-416 and HR-448 of IFP company, hydrogenation catalysts such as ICR154, ICR174, ICR178 and ICR179 of CLG company, hydrogenation catalysts such as HC-P, HC-K UF-210/220 newly developed by UOP company, hydrogenation catalysts such as TK-525, TK-555 and TK-557 of Topsor company, KF-752, KF-756, KF-757, KF-840, KF-848, KF-901, KF-907 and the like hydrogenation catalysts of AKZO corporation.
In the method, the catalyst bed layers of the hydrotreating reactor in the step (a) are arranged into a plurality of layers, preferably 2-8 layers, and a gas dissolving device is arranged between the adjacent catalyst bed layers; the introduced hydrogen is mixed with the reactant flow in the gas dissolving device and then enters the next catalyst bed layer for reaction.
In the above method, one or more catalyst beds, preferably 2 to 8 catalyst beds, may be provided in the make-up hydrogenation reactor. If only one catalyst bed layer is arranged in the supplementary hydrogenation reactor, the liquid-phase hydrogenation reaction material flow is mixed with the refinery gas in the gas dissolver and then enters the top of the supplementary hydrogenation reactor and passes through the catalyst bed layer; if a plurality of catalyst beds are arranged in the supplementary hydrogenation reactor, a gas dissolving device is arranged between the beds, refinery gas and hydrogen are mixed and then enter any gas dissolving device arranged between adjacent catalyst beds, and are mixed with reactant flow from the previous catalyst bed and then enter the next catalyst bed for reaction.
A preferred embodiment is as follows: the catalyst bed layers of the wax oil hydrogenation reactor are arranged into three layers, the catalyst bed layer of the supplementary hydrogenation reactor is arranged into two layers, hydrogen is introduced between the second catalyst and the third catalyst bed layer of the wax oil hydrogenation reactor, and hydrogen and refinery gas are introduced between the catalyst bed layers of the supplementary hydrogenation reactor.
In the method, the wax oil raw oil and the circulating oil are mixed and then enter from the top of the wax oil hydrogenation reactor, the mixed material flow with dissolved hydrogen can pass through the catalyst bed layer from top to bottom in a downward mode, the wax oil raw oil and the circulating oil can also enter from the bottom of the hydrogenation reactor after being mixed, and the mixed material flow with dissolved hydrogen can pass through the catalyst bed layer from bottom to top in an upward mode.
In the method, the mixed material flow of the wax oil hydrogenation reaction effluent dissolved with the refinery gas enters from the top of the supplementary hydrogenation reactor, the mixed material flow dissolved with the refinery gas can pass through the catalyst bed layer from top to bottom, the mixed material flow of the wax oil hydrogenation reaction effluent dissolved with the refinery gas can also enter from the bottom of the supplementary hydrogenation reactor, and the mixed material flow dissolved with the refinery gas can pass through the catalyst bed layer from bottom to top.
In the above method, the previous catalyst bed or the next catalyst bed is based on the flowing direction of the reactant flow, and whether the hydrogenation reaction is an upflow type or a downflow type, the bed in the adjacent beds which is contacted with the reactant flow first is an upper bed and then is a lower bed.
In the method, the refinery gas may comprise one or more of dry gas, liquefied gas and the like. The source of the gas can be one or more of coking, catalytic cracking, thermal cracking, visbreaking and the like.
In the method, if hydrogen and refinery gas are introduced simultaneously in any process, the volume ratio of the introduced hydrogen to the refinery gas is 1: 1-100: 1, preferably 1: 1-50: 1.
In the method, the hydrogenation reaction effluent is separated by a high-pressure separator and/or a low-pressure separator. The high-pressure separator is a conventional gas-liquid separator. The hydrogenation reaction flow is separated in a high-pressure separator to obtain gas and liquid. The low-pressure separator is a conventional gas-liquid separator. The liquid obtained by separation in the high-pressure separator is separated in the high-low pressure separator to obtain gas and liquid.
In the method, the fractionating system used for fractionating comprises a stripping tower and/or a fractionating tower. And the liquid obtained by separation in the low-pressure separator is subjected to steam stripping and/or fractionation in a fractionation system to obtain naphtha, diesel oil and hydrotreated heavy distillate oil.
In the above method, the gas separator used for gas separation is a conventional separator. The gas obtained by separation in the high-pressure separator and the gas obtained by separation in the low-pressure separator are mixed, hydrogen sulfide is removed, then hydrogen, dry gas, liquefied gas and the like are obtained by separation in the gas separator, and if liquid products exist, the gas directly enters a stripping tower and/or a fractionating tower.
In the method, the sulfur content in the diesel oil and/or the hydrotreated heavy fraction is less than 5 mug/g, the nitrogen content is less than 5 mug/g, preferably the sulfur content is less than 3 mug/g, and the nitrogen content is less than 3 mug/g.
In the above method, the hydrocracked raw oil may also include cycle oil used in the full-cycle or partial-cycle operation of the hydrocracking unit liquid, and the cycle oil may include one or more of the hydrocracked liquid products.
In the method, the hydrocracking operation condition is generally that the reaction pressure is 3.0-18.0 MPa, and the volume space velocity of the raw oil is 0.2h-1~6.0h-1The average reaction temperature is 180-450 ℃, and the volume ratio of hydrogen to oil is 300: 1-1500: 1; the preferable operation conditions are that the reaction pressure is 4.0 MPa-16.0 MPa, and the volume airspeed of the raw oil is 0.4h-1~5.0h-1The average reaction temperature is 200-440 ℃, and the volume ratio of hydrogen to oil is 400: 1-1200: 1.
In the above method, the hydrocracking catalyst is a conventional hydrocracking catalyst, and may be a noble metal hydrocracking catalyst or a non-noble metal hydrocracking catalyst. The carrier of the hydrocracking catalyst is alumina and molecular sieve, and the content of the molecular sieve is generally 5wt% -80 wt%. Commercial hydrocracking catalysts are mainly: HC-12, HC-14, HC-24, HC-39, etc. from UOP, FC-12, FC-16, FC-24, 3971, 3976, FC-26, ZHC-02, FC-28, etc. from FRIPP, and ICR126, ICR210, etc. from CHEVRON. Or noble metal catalysts, and the commercial hydrocracking catalysts mainly comprise: HC-28 and HC-35 by Union, and ICR207 and ICR209 by CHEVRON. Conventional hydrocracking pretreatment catalysts and hydrocracking catalysts may also be prepared according to techniques well known in the art.
In the method, the hydrocracking reaction effluent is separated by a high-pressure separator and/or a low-pressure separator. The high-pressure separator is a conventional gas-liquid separator. The hydrogenation reaction flow is separated in a high-pressure separator to obtain gas and liquid. The low-pressure separator is a conventional gas-liquid separator. The gas obtained by separation in the high-pressure separator is recycled after being pressurized by a compressor, and the liquid obtained by separation in the high-pressure separator is separated in the low-pressure separator to obtain the gas and the liquid.
In the method, the hydrocracking fractionation system comprises a fractionating tower. And the liquid obtained by separation in the low-pressure separator is fractionated in a fractionating system to obtain naphtha, aviation kerosene, diesel oil and tail oil.
In the process of gas hydrogenation, the temperature rise of a catalyst bed layer is large due to large reaction heat release, so that the temperature range of the hydrogenation reaction is large, the effect of the hydrogenation reaction is influenced, the generation of carbon deposition of the catalyst is accelerated, and the service cycle of the catalyst is shortened. In the wax oil liquid phase hydrogenation process, hydrogenation reaction is realized through hydrogen dissolved in oil, and the aim of producing clean wax oil products is achieved, but the dissolved hydrogen is excessive and cannot be completely reacted, and the hydrogen dissolved in the hydrogenated oil after the reaction is finished can usually remain 20% -70% of the dissolved hydrogen, so that the hydrogen is ineffectively used, namely, the energy consumption is increased.
According to the invention, by fully utilizing the characteristic that a large amount of hydrogen is still dissolved in oil generated by a wax oil liquid phase circulation hydrogenation process, a supplementary hydrogenation reactor is arranged in the subsequent stage of the wax oil hydrogenation reactor, the refinery gas raw material is dissolved in the wax oil hydrogenation reaction material flow and enters the catalyst bed layer of the supplementary hydrogenation reactor, and the hydrogenation reaction of the gas is completed by utilizing the dissolved hydrogen and the catalyst atmosphere, so that the problem of large gas hydrogenation temperature rise is solved, and the hydrogen dissolved in the wax oil is used for the gas hydrogenation reaction, thereby reducing the hydrogen consumption; or a plurality of catalyst beds are arranged in a further supplementary hydrogenation reactor, part of dry gas or all dry gas raw materials in the mixed gas and wax oil hydrogenation generated oil are mixed to enter the first catalyst bed, and the rest gas and/or hydrogen mixed mixture enters the subsequent catalyst bed. The combined method is generally characterized in that the gas hydrogenation process is completed on the premise of not influencing the quality of the hydrocracking product or further improving the quality of the hydrocracking product to obtain a high-quality hydrocracking product and a gas product, and the two technologies are optimally combined, so that the energy consumption is reduced, the equipment investment is saved, and the operation cost is reduced.
Drawings
FIG. 1 is a flow diagram of a hydrogenation combination process of the present invention.
Wherein: 1-raw oil, 2-raw oil pump, 3-cycle oil, 4-hydrogen dissolver, 5-hydrotreating fresh hydrogen, 5-2-hydrocracking fresh hydrogen, 6-refinery gas raw material, 7-wax oil hydrotreating reactor, 8-vent valve, 9-wax oil hydrotreating reactant stream, 10-hydrotreating high-pressure separator, 11-hydrotreating low-pressure separator, 12-hydrotreating stripping/fractionating system, 13-stripping gas, 14-hydrotreating naphtha, 15-hydrotreating diesel, 16-hydrotreating heavy distillate, 17-hydrotreating high-pressure separator gas, 18-hydrotreating low-pressure separator gas, 19-hydrotreating gas separator, 20-hydrogen, 21-dry gas, 22-liquefied gas, 23-a gas dissolver, 24-a supplementary hydrogenation reactor, 25-a supplementary hydrogenation reaction material flow, 26-a hydrocracking reactor, 27-a hydrocracking reaction material flow, 28-a hydrocracking high-pressure separator, 29-a hydrocracking high-fraction gas, 30-a recycle hydrogen compressor, 31-a hydrocracking high-pressure separator liquid, 32-a hydrocracking low-pressure separator, 33-a hydrocracking low-fraction gas, 34-a hydrocracking low-fraction liquid, 35-a hydrocracking fractionating tower, 36-a hydrocracking naphtha product, 37-hydrocracking aviation kerosene, 38-hydrocracking diesel oil and 39-hydrocracking tail oil.
Detailed Description
The flow and effect of the hydrogenation combination method of the present invention will be further illustrated with reference to the following examples, which should not be construed as limiting the process of the present invention.
The specific implementation mode of the hydrogenation combination method is as follows: raw oil 1 is mixed with cycle oil 3, the mixed material and hydrogen are mixed in a hydrogen dissolver 4 and then enter a wax oil hydrotreating reactor 7, a wax oil hydrotreating reaction material flow 9 and a refinery gas raw material are mixed in a gas dissolver 23 and then enter a supplementary hydrogenation reactor 24, a supplementary hydrotreating reaction material flow 25 enters a hydrotreating high-pressure separator 10 and is separated in the hydrotreating high-pressure separator 10 to obtain hydrotreating high-pressure separator gas 17 and liquid, the liquid enters a hydrotreating low-pressure separator 11 and is separated in the hydrotreating low-pressure separator 11 to obtain hydrotreating low-pressure separator gas 18 and liquid, the liquid and a liquid component obtained by separation of a hydrotreating gas separator 19 are mixed and then enter a hydrotreating stripping/fractionating system 12, and hydrotreating naphtha 14, hydrotreating diesel 15 and hydrotreating heavy distillate 16 are obtained by fractionation under the action of a stripping gas 13 in the fractionating system, the mixture of the hydrotreating high-pressure separator gas 17 and the hydrotreating low-pressure separator gas 18 enters a hydrotreating gas separator 19, and hydrogen, dry gas and liquefied gas products are obtained by separation in the hydrotreating gas separator 19. The cycle oil 3 can be obtained directly from the wax oil hydrotreating reaction material flow 9, and also can be obtained from the liquid obtained by the separation in the hydrotreating high-pressure separator 10. The method comprises the steps of mixing hydrotreated diesel 15, hydrotreated heavy fraction oil 16 and recycle hydrogen, enabling the mixture to enter a hydrocracking reactor 26, enabling the mixture to pass through a hydrocracking catalyst bed layer, carrying out gas-liquid separation on a hydrocracking reaction material flow 27 in a hydrocracking high-pressure separator 28, boosting a separated hydrocracking high-fraction gas 29 through a recycle hydrogen compressor 30, recycling the gas-liquid separation, enabling a separated hydrocracking high-pressure separator liquid 31 to continue to carry out gas-liquid separation in a hydrocracking low-pressure separator 32, separating hydrocracking low-fraction gas 33, and enabling a separated hydrocracking low-fraction liquid 34 to continue to enter a hydrocracking fractionating tower 35 for fractionation to obtain a hydrocracking naphtha product 36, hydrocracking aviation kerosene 37, hydrocracking diesel 38 and hydrocracking tail oil 39.
The following examples further illustrate specific aspects of the present invention. Experimental studies were conducted using FF-56 hydrotreating catalyst and FC-32 cracking catalyst developed and produced by FRIPP development.
TABLE 1 wax oil feedstock Primary Properties
| Wax oil feedstock | Raw oil 1 | |
| Density, g/cm3 | 0.913 | 0.924 |
| Range of distillation range, deg.C | 320~547 | 335~560 |
| Sulfur content, wt.% | 1.32 | 2.16 |
| Nitrogen content, wt% | 0.146 | 0.089 |
TABLE 2 refinery gas feedstock key properties
| Refinery gas feedstock | Dry gas | Liquefied gas | Mixed gas |
| Gas composition | |||
| H2 | 7.0 | 0 | 3.5 |
| CH4 | 12.6 | 0 | 2.9 |
| C2H6 | 55.3 | 0 | 27.1 |
| C2H4 | 5.6 | 0.1 | 4.6 |
| C3 H8 | 10.8 | 16.0 | 13.6 |
| C3 H6 | 2.7 | 6.5 | 4.5 |
| C3 H4 | 0 | 0 | 0 |
| C4 H10 | 5.3 | 34.5 | 20.5 |
| C4 H8 | 0.5 | 33.1 | 19.1 |
| C4 H6 | 0 | 1.2 | 0.5 |
| C5 + | 0.1 | 8.6 | 3.6 |
| CO | 0.005 | 0 | 0.002 |
| CO2 | 0.01 | 0 | 0.008 |
Table 3 examples process conditions and main product properties
| Process conditions | Example 1 | Example 2 | Example 3 | Example 4 | Example 5 |
| Hydrogenation reactor operating conditions | |||||
| Raw oil | Raw oil 1 | Raw oil 1 | Raw oil 1 | Raw oil 2 | Raw oil 2 |
| Operating conditions of wax oil hydrogenation reactor | |||||
| Reaction pressure, MPa | 15.0 | 13.0 | 13.0 | 12.0 | 17.0 |
| Average reaction temperature,. degree.C | 380 | 370 | 370 | 375 | 385 |
| Volume space velocity of fresh raw oil, h-1 | 1.2 | 1.0 | 1.0 | 0.8 | 1.5 |
| Circulation ratio | 2:1 | 2.5:1 | 2.5:1 | 4:1 | 3:1 |
| Make-up of hydrogenation reactor operating conditions | |||||
| Refinery gas feedstock at reactor inlet | Dry gas | Mixed gas | Dry gas | Liquefied gas | Mixed gas |
| Reaction pressure, MPa | 15.0 | 13.0 | 13.0 | 12.0 | 17.0 |
| Average reaction temperature,. degree.C | 380 | 370 | 370 | 375 | 385 |
| Volume space velocity of fresh raw oil, h-1 | 20.0 | 15.0 | 15.0 | 10 | 16 |
| Two-bed inlet refinery gas raw material | — | — | Liquefied gas | — | — |
| Volume ratio of hydrogen dissolved in inlet of two-bed layer to raw material of refinery gas | — | — | 95:5 | — | — |
| Hydrocracking process conditions | |||||
| Raw oil | Hydrotreating heavy distillate | Hydrotreating heavy distillate | Hydrotreating heavy distillate | Hydrotreating heavy distillate | Hydrotreating heavy distillate |
| Reaction pressure, MPa | 10.0 | 13.0 | 13.0 | 10.0 | 15.0 |
| Volumetric space velocity h-1 | 1.2 | 2.0 | 2.0 | 0.8 | 2.5 |
| Average reaction temperature,. degree.C | 330 | 340 | 340 | 325 | 345 |
| Volume ratio of hydrogen to oil | 1000 | 1200 | 1200 | 800 | 1000 |
| Dry gas product | |||||
| Olefin content, v% | 0 | 0 | 0 | 0 | 0 |
| Liquefied gas product | |||||
| Olefin content, v% | 0 | 0 | 0 | 0 | 0 |
| CO+CO2，µg/g | 0 | 0 | 0 | 0 | 0 |
| Naphtha product | |||||
| Sulphur content, μ g/g | 0.3 | 0.4 | 0.4 | 0.3 | 0.4 |
| Aviation kerosene product | |||||
| Smoke point, mm | 27 | 26 | 26 | 25 | 29 |
| Diesel oil product | |||||
| Density, g/cm3 | 0.813 | 0.821 | 0.821 | 0.830 | 0.825 |
| Sulphur content, μ g/g | 2 | 5 | 5 | 3 | 4 |
| Cetane number | 60 | 57 | 57 | 53 | 56 |
| Tail oil | |||||
| BMCI | 4.6 | 6.4 | 6.4 | 8.9 | 7.5 |
| Sulphur content, μ g/g | 0.5 | 0.6 | 0.6 | 0.3 | 0.2 |
As can be seen from the examples, wax oil feedstock and refinery gas feedstock can be directly used to produce high quality hydrocracked products and clean gas products by the hydrocombination process of the present technology.
Claims (18)
1. The combined processing method of refinery gas is characterized by comprising the following steps: the method comprises the following steps:
(a) mixing the wax oil raw oil and the circulating oil with hydrogen in hydrogen dissolving equipment, and then entering a hydrotreating catalyst bed layer in a wax oil hydrotreating reactor to react under the liquid-phase hydrogenation operation condition;
(b) mixing the reactant flow obtained in the step (a) with refinery gas and hydrogen in a gas dissolving device, and then allowing the mixture to enter a hydrogenation catalyst bed layer in a supplementary hydrogenation reactor to react under the liquid-phase hydrogenation operation condition;
(c) separating the hydrogenation reaction effluent in the step (b) into a gas phase and a liquid phase, continuously separating the gas phase obtained by separation after removing hydrogen sulfide to obtain hydrogen and hydrotreated refinery gas, fractionating the liquid phase obtained by separation to obtain naphtha, diesel oil and hydrotreated heavy distillate oil, and returning part of the hydrogenation reaction effluent obtained in the step (a) and/or part of the hydrogenation reaction material flow obtained in the step (b) and/or part of the liquid phase obtained by separation of a high-pressure separator as circulating oil to hydrogen dissolving equipment;
(d) mixing part or all of the diesel oil and/or the hydrotreated heavy fraction obtained in the step (c) with hydrogen, allowing the mixture to enter a hydrocracking catalyst bed layer in a hydrocracking reactor for reaction, separating reactant streams in a high-pressure separator, recycling the separated gas, and fractionating the separated liquid in a fractionating tower to obtain naphtha, aviation kerosene, diesel oil and tail oil.
2. The method of claim 1, wherein: the wax oil raw material oil is one or more of VGO, CGO, HGO, HLCO, DAO, coal tar and coal liquefaction oil.
3. The method of claim 1, wherein: the wax oil hydrotreating operation conditions are that the reaction pressure is 3.0 MPa-20.0 MPa, and the volume space velocity of the wax oil raw material oil is 0.2h-1~8.0h-1The average reaction temperature is 180-450 ℃, and the ratio of the circulating oil to the wax oil raw oil is 0.5: 1-10: 1.
4. The method of claim 3, wherein: the wax oil hydrotreating operation conditions are that the reaction pressure is 4.0 MPa-18.0 MPa, and the volume space velocity of the wax oil raw material oil is 0.5h-1~6.0h-1The average reaction temperature is 200-440 ℃, and the ratio of the circulating oil to the wax oil raw oil is 0.6: 1-8: 1.
5. The method of claim 1, wherein: the supplementary hydrotreating operation conditions are that the reaction pressure is 3.0 MPa-20.0 MPa, and the volume airspeed of the wax oil raw material oil is 0.5h-1~40.0h-1The average reaction temperature is 180-450 ℃.
6. The method of claim 5, wherein: the supplementary hydrotreating operation condition is that the reaction pressure is 4.0 MPa-18.0 MPaThe volume space velocity of the wax oil raw oil is 0.8h-1~30.0h-1The average reaction temperature is 200-440 ℃.
7. The method of claim 1, wherein: the hydrogenation active components of the hydrogenation catalyst used in the wax oil hydrogenation reactor and the hydrogenation active components of the hydrogenation catalyst used in the supplementary hydrogenation reactor are one or more of Co, Mo, W and Ni, the weight content of the hydrogenation active components is 5-70 percent calculated by oxides, and the carrier of the hydrogenation catalyst is one or more of alumina, amorphous silicon aluminum, silicon oxide and titanium oxide.
8. The method of claim 1, wherein: 1-8 catalyst beds are arranged in the wax oil hydrotreating reactor and/or the supplementary hydrotreating reactor.
9. The method of claim 8, wherein: three catalyst beds are arranged in the wax oil hydrogenation reactor, two catalyst beds are arranged in the supplementary hydrogenation reactor, gas dissolving equipment is arranged between the beds, hydrogen enters the gas dissolving equipment between the beds of the wax oil hydrogenation reactor, and hydrogen and refinery gas enter the gas dissolving equipment between the beds of the supplementary hydrogenation reactor.
10. The method of claim 1, wherein: the refinery gas is one or more of dry gas and liquefied gas, and the gas is one or more of coking, catalytic cracking and thermal cracking reaction.
11. The method according to claim 1 or 9, characterized in that: when hydrogen and refinery gas are introduced simultaneously, the volume ratio of the introduced hydrogen to the refinery gas is 1: 1-100: 1.
12. The method of claim 1, wherein: and (d) in the step (d), the sulfur content in the diesel oil and/or the hydrotreated heavy fraction is less than 5 mug/g, and the nitrogen content is less than 5 mug/g.
13. The method of claim 12, wherein: and (d) in the step (d), the sulfur content in the diesel oil and/or the hydrotreated heavy fraction is less than 3 mug/g, and the nitrogen content is less than 3 mug/g.
14. The method of claim 1, wherein: the hydrocracking operation conditions are that the reaction pressure is 3.0MPa to 18.0MPa, and the volume airspeed of the raw oil is 0.2h-1~6.0h-1The average reaction temperature is 180-450 ℃, and the volume ratio of hydrogen to oil is 300: 1-1500: 1.
15. The method of claim 14, wherein: the hydrocracking operation conditions are that the reaction pressure is 4.0MPa to 16.0MPa, and the volume airspeed of the raw oil is 0.4h-1~5.0h-1The average reaction temperature is 200-440 ℃, and the volume ratio of hydrogen to oil is 400: 1-1200: 1.
16. The method of claim 1, wherein: the hydrocracking catalyst is a noble metal hydrocracking catalyst or a non-noble metal hydrocracking catalyst, the carrier of the hydrocracking catalyst is alumina and a molecular sieve, and the content of the molecular sieve is 5wt% -80 wt%.
17. The method of claim 1, wherein: and separating the hydrocracking reaction material flow in a high-pressure separator to obtain gas and liquid, pressurizing the gas obtained by separation in the high-pressure separator by a compressor and then recycling the gas, and separating the liquid obtained by separation in the high-pressure separator in a low-pressure separator to obtain the gas and the liquid.
18. The method of claim 17, wherein: and (4) fractionating the liquid obtained by separation in the low-pressure separator in a fractionating system to obtain naphtha, aviation kerosene, diesel oil and tail oil.
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| CN101880549A (en) * | 2009-05-06 | 2010-11-10 | 中国石油化工股份有限公司 | Hydrogenation method for liquefied gas fraction |
| CN102268293A (en) * | 2010-06-07 | 2011-12-07 | 中国石油化工集团公司 | Hydrogenation combined process |
| CN104560132A (en) * | 2013-10-29 | 2015-04-29 | 中国石油化工股份有限公司 | Method for hydrotreatment of continuous liquid-phase wax oil |
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| CN101880549A (en) * | 2009-05-06 | 2010-11-10 | 中国石油化工股份有限公司 | Hydrogenation method for liquefied gas fraction |
| CN102268293A (en) * | 2010-06-07 | 2011-12-07 | 中国石油化工集团公司 | Hydrogenation combined process |
| CN104560132A (en) * | 2013-10-29 | 2015-04-29 | 中国石油化工股份有限公司 | Method for hydrotreatment of continuous liquid-phase wax oil |
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