A method for hydrogenating and dearomatizing distillate oil and a method for preparing white oil

Abstract

The present invention provides a method for hydrogenating and dearomatizing distillate oil and a method for preparing white oil. Based on the method of the present invention, high-quality middle distillate oil with a low aromatic content can be directly produced from distillate oil. The method comprises the following steps: reacting the distillate oil with a hydrogen-containing stream in a first hydrogenation reactor to obtain a first reaction product, wherein the conversion rate of the fraction with an initial boiling point greater than 400°C is not greater than 40% and the isoparaffin content in the first reaction product is not less than 50% by weight; separating and fractionating the first reaction product to obtain a first middle distillate oil with a boiling range of 160 to 400°C and a first heavy oil with an initial boiling point greater than 400°C; reacting the first heavy oil with a hydrogen-containing stream in a second hydrogenation reactor to obtain a second reaction product, wherein the conversion rate of the fraction with an initial boiling point greater than 400°C is not less than 40%, preferably not less than 65%; and separating and fractionating the second reaction product.

Description

Technical Field

[0001] The invention belongs to the technical field of petrochemical industry, and in particular relates to a method for hydrogenating and dearomatizing distillate oil and a method for preparing white oil. Background Art

[0002] The aromatic content in oil products is one of the important indicators. However, the hydrogenation reaction of aromatics is difficult to carry out at high temperatures due to the limitations of thermodynamic equilibrium. Therefore, when preparing low-aromatic product oil by hydrogenation, there are often problems such as the need for special catalysts or specific specifications of raw oil. In addition, the product oil can only meet the aromatic content specification requirements in diesel, and is far from meeting the aromatic content requirements of high-value-added white oil, which seriously affects the economic benefits of the hydrogenation process.

[0003] Patent document CN 1119395C discloses a two-stage hydrodearomatization method for distillate oil. The distillate feedstock enters a first reactor, where it reacts in the presence of hydrogen and a non-precious metal catalyst. The resulting material enters a hot high-pressure separator. The material exiting the separator's bottom enters a second reactor, where it reacts in the presence of hydrogen and a precious metal catalyst. The resulting material is separated in a high-pressure separator and then a low-pressure separator to produce a liquid product. The hydrogen-rich gas at the top of the high-pressure separator is recycled. This method can reduce the aromatics content of the product to below 10% by weight, with a yield of over 95% by weight.

[0004] Patent document CN1173012C discloses a method for deep desulfurization and dearomatization of diesel. The raw oil is mixed with hydrogen and enters a first reactor, where it contacts a hydroreforming catalyst, a hydrocracking catalyst, or a hydrorefining catalyst. The reaction effluent is stripped in a hydrogen stripper under high temperature and high pressure and then mixed with hydrogen and enters a second reactor where it contacts a conventional hydrorefining catalyst. The effluent from the second reactor enters a high-pressure separator, a low-pressure separator, and a fractionating tower in sequence. The fractionating tower separates the reaction products into naphtha and diesel fractions. The hydrogen-rich gas stream separated from the high-pressure separator and the high-pressure stripping tower is mixed with fresh hydrogen and divided into two parts and sent to the first and second reactors, respectively.

[0005] The method disclosed in patent document CN 1119395C is primarily targeted at diesel fraction feedstocks and is not suitable for processing feedstocks with higher distillation ranges, such as Fischer-Tropsch refining tail oil, cracking tail oil, and Fischer-Tropsch full-fraction oil. The patent requires the use of a precious metal catalyst, which is susceptible to poisoning and is not conducive to long-term operation of the device. Furthermore, the patent can only reduce the aromatic content in the product oil to below 10%, and cannot further reduce the aromatic content in the distillate oil, far from meeting the aromatic content requirements of white oil.

[0006] Patent document CN1173012C discloses a method for deep desulfurization and dearomatization of diesel, which is only suitable for treating diesel with a distillation range of 180-390°C and is not suitable for treating raw oils with higher distillation ranges, such as Fischer-Tropsch refined tail oil, cracking tail oil, and Fischer-Tropsch full-fraction oil. Moreover, the product oil produced by this patent can only meet the specification requirements for aromatic content in diesel, and is far from meeting the requirements for aromatic content in white oil. Summary of the Invention

[0007] The present invention provides a method for hydrogenating and dearomatizing distillate oil. Based on the method of the present invention, distillate oil can be used to directly produce high-quality middle distillate oil with a low aromatic content. The obtained middle distillate oil can be finely cut to produce white oil with an aromatic content that meets the standard.

[0008] To achieve the purpose of the present invention, the present invention provides the following technical solutions:

[0009] The present invention provides a method for hydrodearomatization of distillate oil, comprising the following steps:

[0010] reacting a distillate oil as a raw material with a hydrogen-containing stream in a first hydrogenation reactor in the presence of a first hydroisomerization catalyst to obtain a first reaction product, and controlling the conversion depth and isomerization depth during the reaction so that the conversion rate of the fraction having an initial boiling point greater than 400° C. is no greater than 40% and the isoparaffin content in the first reaction product is no less than 50% by weight;

[0011] Separating and fractionating the first reaction product to obtain a first middle distillate oil with a distillation range of 160 to 400° C. and a first heavy oil with an initial boiling point greater than 400° C.;

[0012] reacting the first heavy oil with a hydrogen-containing stream in a second hydrogenation reactor in the presence of a second hydroisomerization catalyst to obtain a second reaction product, and controlling the conversion depth during the reaction so that the conversion rate of the fraction having an initial boiling point greater than 400° C. is not less than 40%, preferably not less than 65%;

[0013] The second reaction product is separated and fractionated to obtain a second middle distillate oil with a distillation range of 160-400° C. and a second heavy oil with an initial distillation point greater than 400° C.

[0014] In some embodiments, the second heavy oil is circulated to the second hydrogenation reactor to participate in the reaction.

[0015] In some embodiments, the first reaction product and the second reaction product are separated and fractionated in separate devices, respectively, or the first reaction product and the second reaction product are separated and fractionated in the same device.

[0016] In some embodiments, the first hydroisomerization cracking catalyst is a hydroisomerization cracking catalyst containing amorphous silica-alumina;

[0017] Preferably, in the first hydroisomerization cracking catalyst, based on the total weight of the catalyst, the content of amorphous silica-alumina is 32-64.5%, the content of aluminum oxide is 20-50%, the content of tungsten oxide is 5-30%, and the content of nickel oxide is 0.5-9%; preferably, the pore volume of the amorphous silica-alumina used to prepare the first hydroisomerization cracking catalyst is 1.0-1.6 mL / g, and the specific surface area is 300-500 m 2 / g.

[0018] In some embodiments, the second hydroisomerization cracking catalyst is a hydroisomerization cracking catalyst containing a molecular sieve.

[0019] In some embodiments, the second hydrogenation reactor is sequentially loaded with the second hydroisomerization cracking catalyst and the hydrorefining catalyst along the flow direction of the reaction materials;

[0020] Preferably, in the second hydrogenation reactor, the loading volume ratio of the second hydroisomerization cracking catalyst to the hydrorefining catalyst is 100:10-30.

[0021] In some embodiments, the hydrotreating catalyst includes a carrier and an active component supported on the carrier; based on the weight of the hydrotreating catalyst, the active component includes 10wt% to 50wt% of tungsten oxide, 5wt% to 30wt% of molybdenum oxide, and 0.5wt% to 20wt% of nickel oxide; preferably, the carrier includes one or both of aluminum oxide and silicon oxide.

[0022] In some embodiments, the specific surface area of ​​the hydrotreating catalyst is 100 to 400 m 2 / g, the pore volume is 0.3~1.2ml / g, the infrared acid content is 0.3~1.0mmol / g, and the ratio of B acid / L acid is 1.5~2.0.

[0023] In some embodiments, the distillate oil used as the raw material has a distillation range of 180-800°C.

[0024] In some embodiments, the distillate oil used as the raw material includes a mixed oil of one or more of synthetic light oil, heavy oil, synthetic wax full fraction, Fischer-Tropsch refined tail oil, and Fischer-Tropsch cracking tail oil from a Fischer-Tropsch synthesis process.

[0025] In some embodiments, the reaction conditions of the first hydrogenation reactor include: reaction temperature 310-370 ° C, volume space velocity 1-4h -1 , hydrogen-oil volume ratio 500-1500:1, hydrogen partial pressure: 5-15MPa;

[0026] The reaction conditions of the second hydrogenation reactor include: reaction temperature 320-375 ° C, volume space velocity 1-2h -1 , hydrogen-oil volume ratio 500-1500:1, hydrogen partial pressure: 5-15MPa.

[0027] The present invention also provides a method for preparing white oil, comprising fractionating the first middle distillate oil and / or the second middle distillate oil obtained by the method described above to obtain the white oil.

[0028] The technical solution provided by the present invention has the following beneficial effects:

[0029] The present invention can directly produce high-quality middle distillate oil with low aromatic content from distillate oil. Finely cutting the middle distillate oil can produce different grades of white oil (e.g., light white oil, industrial white oil, cosmetic-grade white oil, and food-grade white oil) with aromatic content that meets standards. In some embodiments, the total yield of the middle distillate oil is greater than 76%.

[0030] The distillate oil hydrodearomatization method provided by the present invention is applicable to raw oils with a wider distillation range and heavier weight, such as Fischer-Tropsch refining tail oil and cracking tail oil.

[0031] The distillate oil hydrodearomatization method provided by the present invention can be implemented with relatively few changes to the existing hydrogenation unit process, greatly saving the investment and time for the transformation of the unit, and facilitating the rapid transformation and upgrading of the existing unit to produce high-value-added white oil products.

[0032] In the distillate oil hydrodearomatization method provided by the present invention, the first and second hydrogenation reactors are independently arranged, and catalysts with different requirements can be flexibly loaded. The reaction temperature and volumetric space velocity can also be independently controlled, thereby easily controlling the conversion rate and isomerization depth of the fraction with an initial boiling point greater than 400°C, conveniently adjusting the yield of white oil of different grades, and providing flexible product solutions.

[0033] In the present invention, the product obtained from the first hydrogenation reactor is first separated and fractionated to promptly separate the middle distillate oil at 160-400°C, thereby avoiding multiple cracking and over-cracking of this fraction in the second hydrogenation reactor. On the one hand, this improves the yield of the middle distillate, and on the other hand, significantly reduces the processing capacity of the second hydrogenation reactor, thereby reducing equipment investment and saving operating costs. DETAILED DESCRIPTION

[0034] In order to facilitate the understanding of the present invention, the present invention will be further described below in conjunction with examples. It should be understood that the following examples are only for a better understanding of the present invention and do not mean that the present invention is limited to the following examples.

[0035] Unless otherwise defined, all technical and scientific terms used herein have the same meaning as those generally understood by those skilled in the art to which the present invention belongs. In addition, the terms "first," "second," "third," etc. are used only for descriptive purposes and are not to be construed as indicating or suggesting relative importance.

[0036] Where specific experimental steps or conditions are not specified in the examples, the corresponding conventional experimental steps or conditions in the art can be used. Reagents or instruments used without specifying the manufacturer are all commercially available conventional products.

[0037] The present invention provides a method for hydrogenating and dearomatizing distillate oil, which mainly comprises the following steps: reacting distillate oil (i.e., feed oil) as a raw material with a hydrogen-containing stream in the presence of a first hydroisomerization cracking catalyst in a first hydrogenation reactor to obtain a first reaction product, and controlling the conversion depth and isomerization depth during the reaction so that the conversion rate of the fraction with an initial boiling point greater than 400°C is not greater than 40% and the isoparaffin content in the first reaction product is not less than 50% by weight; separating and fractionating the first reaction product to obtain gas, naphtha, and a distillation range of 160°C; The invention relates to a method for preparing a first middle distillate oil with an initial boiling point of greater than 400°C and a first heavy oil with an initial boiling point of greater than 400°C; reacting the first heavy oil with a hydrogen-containing stream in a second hydrogenation reactor in the presence of a second hydroisomerization catalyst to obtain a second reaction product, and controlling the conversion depth during the reaction so that the conversion rate of the fraction with an initial boiling point greater than 400°C is not less than 40%, and more preferably not less than 65% to improve reaction economy; and separating and fractionating the second reaction product to obtain a second middle distillate oil with a distillation range of 160-400°C and a second heavy oil with an initial boiling point greater than 400°C.

[0038] The inventors have discovered that when a feedstock is hydrocracking to produce a low-aromatic product oil, the aromatic content in the product oil is affected not only by the aromatic content of the feedstock, but also by the hydrotreating process. During the hydrotreating process, dearomatization occurs, and hydrocracking intermediates are converted into aromatics. This is particularly true during multiple cracking processes, which are more likely to generate aromatics, resulting in substandard aromatics in the product. When the volumetric space velocity is too low and / or the reaction temperature is too high and / or the hydrogen partial pressure is insufficient in the reaction process conditions, the conversion of intermediates into aromatics is exacerbated, resulting in an inability to deeply remove aromatics from the product oil, and the technical requirements for aromatic content in white oil cannot be met. The present invention unexpectedly discovered that by subjecting the distillate oil used as a feedstock to graded hydrogenation treatment and limiting the conversion and / or isomerization cracking depth of the first and second hydrocracking processes, a low-aromatic middle distillate oil can be directly produced from the feedstock oil (e.g., a distillate oil with a boiling range of 180-800°C). The middle distillate oil can be finely cut to obtain high-quality light white oil, industrial white oil, cosmetic-grade white oil, and food-grade white oil with a qualified aromatic content, and the middle distillate oil yield is high. The control of the conversion depth and isomerization depth in the first hydrogenation reactor, as well as the conversion depth in the second hydrogenation reactor, can be achieved by controlling the reaction temperature and volumetric space velocity; specifically, in the present invention, the conversion rate of the fraction with an initial boiling point greater than 400°C is mainly achieved by controlling the volumetric space velocity and reaction temperature; in the actual operation of hydrocracking, the volumetric space velocity and reaction temperature are commonly used effective control means, such as increasing the temperature to increase the conversion rate, and decreasing the volumetric space velocity to increase the conversion rate.

[0039] In some embodiments, the second heavy oil obtained by separating and fractionating the second reaction product is recycled to the second hydrogenation reactor to continue the reaction.

[0040] In some embodiments, the first reaction product and the second reaction product are separated and fractionated in separate devices respectively; or the first reaction product and the second reaction product are separated and fractionated in the same set of devices, that is, the second reaction product is returned to the device for separating and fractionating the first reaction product for separation and fractionation, and the second middle distillate oil will be obtained together with the first middle distillate oil, and the second heavy oil will be obtained together with the first heavy oil. The specific separation and fractionation operations for separating and fractionating the reaction products in the first and second hydrogenation reactors belong to conventional techniques in the field and are not particularly limited. Conventional separation and fractionation operations in the field can be used for the separation and fractionation. For example: in some embodiments, the product obtained in the hydrogenation reactor is separated by a hot high-pressure separator to obtain liquid phase 1 and gas phase 1, gas phase 1 is sent to a cold high-pressure separator to obtain liquid phase 2 and gas phase 2, liquid phase 1 is sent to a hot low-pressure separator to obtain gas phase 3 (light hydrocarbons and dissolved hydrogen) and liquid phase 3 (product oil), gas phase 2 and gas phase 3 are mixed and sent to the downstream section, and liquid phase 2 and liquid phase 3 are mixed and sent to a normal vacuum distillation tower for fractionation to obtain light hydrocarbons (<C5 +), naphtha, middle distillate (160-400°C fraction), and heavy oil (initial boiling point greater than 400°C), the above separation operations are conventional techniques in the art. In some embodiments, the naphtha obtained by separating and fractionating the first reaction product and the second reaction product is a C5-160°C fraction.

[0041] The distillate oil hydrodearomatization method provided by the present invention utilizes a two-stage hydrotreating process and limits the first and second hydrocracking steps to specific reaction levels. This method directly produces middle distillates with low aromatic content without the need for specialized hydrogenation catalysts. This method enables the production of middle distillates with low aromatic content with minimal modifications to existing hydrogenation process flow, significantly reducing equipment modification investment and time, and facilitating the rapid modification and upgrading of existing equipment to produce high-value-added white oil products.

[0042] In the distillate oil hydrodearomatization method of the present invention, the first and second hydrogenation reactors are independently provided, so that the required type and ratio of catalyst can be loaded into each hydrogenation reactor according to reaction requirements. The reaction temperature and volumetric space velocity can also be independently controlled, and the two-stage hydrogenation treatment can be easily controlled separately. The conversion rate and isomerization depth of the fraction with an initial boiling point greater than 400° C. can be easily controlled, and the yield of white oil of different brands can be conveniently adjusted, and the product plan is flexible.

[0043] In the distillate oil hydrodearomatization method of the present invention, the first reaction product obtained by the reaction in the first hydrogenation reactor is first subjected to separation and fractionation, and the middle distillate oil with a temperature of 160-400° C. is separated out in time, thereby avoiding multiple cracking and over-cracking of the distillate segment in the second hydrogenation reactor. On the one hand, the yield of the middle distillate oil is improved, and on the other hand, the processing capacity of the second hydrogenation reactor is significantly reduced, thereby reducing equipment investment and saving operating costs.

[0044] The first hydroisomerization cracking catalyst required for the reaction in the first hydrogenation reactor can adopt the hydroisomerization cracking catalyst that can realize the corresponding function available in the art. As a preferred embodiment, the first hydroisomerization cracking catalyst can adopt a hydroisomerization cracking catalyst containing amorphous silica-alumina, such as the hydrocracking catalyst disclosed in CN110433819A. The inventors have found that the use of a hydroisomerization cracking catalyst based on amorphous silica-alumina in the first hydrogenation reactor is beneficial to the diffusion of Fischer-Tropsch synthesis macromolecules, and the use of molecular sieves containing a large amount of strong acid and micropores can reduce the secondary cracking of the raw material, thereby reducing the yield of light hydrocarbons and naphtha, while avoiding the generation of aromatics in multiple cracking processes, which helps to control the aromatic content in the product oil. In some preferred embodiments, the first hydroisomerization cracking catalyst has, based on the total weight of the catalyst, amorphous silica-alumina content of 32-64.5%, aluminum oxide content of 20-50%, tungsten oxide content of 5-30%, and nickel oxide content of 0.5-9%; the amorphous silica-alumina used to prepare the catalyst has a pore volume of 1.0-1.6 mL / g and a specific surface area of ​​300-500 m 2 / g; the alumina in the catalyst is derived from the use of an aluminum-containing binder, such as a pseudo-boehmite binder; the catalyst can be prepared by referring to the method in Chinese patent application CN110433819A, and the specific preparation process is not described here.

[0045] The second hydroisomerization cracking catalyst required for the reaction in the second hydrogenation reactor of the present invention can adopt a hydroisomerization cracking catalyst that can achieve the corresponding function in the art. In some preferred embodiments, the second hydroisomerization cracking catalyst is a hydroisomerization cracking catalyst containing a molecular sieve, and can adopt a hydroisomerization cracking catalyst that can achieve the corresponding function in the art containing a molecular sieve. Such catalysts can be obtained through commercial channels, such as but not limited to RT-5 developed by the Petrochemical Science Research Institute; FC16, FC14, FC12, H-06, 3762, 3281 developed by Fushun Petrochemical Research Institute; HC-11, HC-14, HC-16, HC-18, HC-24, HC-28, HC-22, HC-33 developed by Union; ICR-117, ICR-126, ICR-220 developed by Chevron; DHC-39, DHC-41 developed by UOP, etc.

[0046] In some embodiments, the second hydroisomerization cracking catalyst and the hydrorefining catalyst are sequentially loaded into the second hydrogenation reactor along the flow direction of the reactants. In some preferred embodiments, the loading volume ratio of the second hydroisomerization cracking catalyst to the hydrorefining catalyst in the second hydrogenation reactor is 100:10-30.

[0047] Regarding the case where the second hydrogenation reactor uses a hydrorefining catalyst, the hydrorefining catalyst used can be a corresponding hydrorefining catalyst conventional in the art. In some preferred embodiments, the hydrorefining catalyst includes a carrier and an active component loaded on the carrier; based on the weight of the hydrorefining catalyst, the active component includes 10wt% to 50wt% of tungsten oxide, 5wt% to 30wt% of molybdenum oxide, and 0.5wt% to 20wt% of nickel oxide. In some embodiments, the balance is the carrier; preferably, the carrier includes one or two of aluminum oxide and silicon oxide, such as aluminum oxide or a mixture of aluminum oxide and silicon oxide; further preferably, the specific surface area of ​​the hydrorefining catalyst is 100 to 400m 2 / g, pore volume of 0.3-1.2 ml / g, infrared acid content of 0.3-1.0 mmol / g, and B acid / L acid ratio of 1.5-2.0; the hydrorefining catalyst can be prepared by a conventional catalyst preparation method, for example, by the following conventional method: first preparing a carrier, then impregnating it with an equal volume of a water-soluble salt solution of an active metal element, and then drying and calcining to obtain the catalyst. As an example, it is prepared by the following method: 45-65 parts (the following parts are all by weight) of pseudo-boehmite, 25-65 parts of amorphous silica-alumina, 2-10 parts of sesbania powder, 5-15 parts of nitric acid, etc. are mixed (for example, mixed for 0.5-3 hours) and then extruded into strips through an extruder, dried at 100-120°C (for example, dried for 4 hours), calcined at 600-700°C for 4-8 hours (for example, calcined at 650°C for 6 hours), and treated at 400-500°C for 2-6 hours (for example, 450°C for 4 hours) to obtain carrier A, and the mass ratio of water vapor to carrier A is 0.1-0.5:1; carrier A is made of a mixture containing a solubility of 0 .2-12wt% aluminum nitrate in acetic acid solution for 2-8h, the solution pH value is 2.5-4.5, filtered and washed with water until the washing liquid is neutral, then dried at 100-120℃ (for example, dried for 4 hours), calcined at 600-700℃ for 4-8 hours (for example, calcined at 650℃ for 6 hours) to obtain carrier B; an equal volume of water-soluble salt solution of active metal elements (for example, an anion water-soluble salt solution containing tungsten, molybdenum, and nickel) is impregnated, dried at 120℃, and calcined at 600-700℃ for 4-8 hours (for example, calcined at 650℃ for 6 hours) to obtain a catalyst; those skilled in the art can adjust and determine the corresponding raw materials and amounts according to the needs of the catalyst composition. The hydrorefining catalysts used in subsequent Examples 1 and 4 were all prepared according to this exemplary method, the main difference being that amorphous silicon aluminum was not added. Phosphoric acid was also added to the hydrorefining catalyst of Example 1 when preparing carrier A, which will not be repeated here.

[0048] The hydrorefining catalyst may also be a commercially available catalyst that meets the above requirements.

[0049] The method for hydrodearomatization of distillate oil provided by the present invention can be applicable to raw oils with a wider distillation range and heavier weight, for example, it can be applicable to distillate oils with a distillation range of 180-800°C as raw oils. Specifically, for example, the distillate oils used as raw materials include one or more mixed oils of synthetic light oil (distillation range 180-400°C) from the Fischer-Tropsch synthesis process, heavy oil (distillation range greater than 400°C), full fractions of synthetic wax, Fischer-Tropsch refined tail oil, and Fischer-Tropsch cracking tail oil.

[0050] The reaction conditions in the first hydrogenation reactor and the second hydrogenation reactor of the present invention are based on the control requirements of the present invention for the corresponding conversion depth and isomerization depth during the reaction process. Those skilled in the art can make corresponding adjustments and determinations for the specific reaction conditions based on this goal, that is, in the first hydrogenation reactor, the conversion depth and isomerization depth are controlled during the reaction process so that the conversion rate of the fraction with an initial boiling point greater than 400°C is not greater than 40% and the isoparaffin content in the first reaction product is not less than 50% by weight, and in the second hydrogenation reactor, the conversion depth is controlled during the reaction process so that the conversion rate of the fraction with an initial boiling point greater than 400°C is not less than 40%. In some embodiments, in order to control the low conversion rate of the first hydrogenation reactor, a relatively large volume space velocity is used, and the device processing capacity is large; a relatively low reaction temperature can also be used to extend the catalyst life. In some embodiments, the reaction conditions of the first hydrogenation reactor include: a reaction temperature of 310-370°C, a volume space velocity of 1-4h -1 , hydrogen-oil volume ratio 500-1500:1, hydrogen partial pressure: 5-15MPa; reaction conditions of the second hydrogenation reactor include: reaction temperature 320-375℃, volume space velocity 1-2h -1 , hydrogen-to-oil volume ratio 500-1500:1, hydrogen partial pressure: 5-15MPa, within the above conditions, the control of specific conversion depth and isomerization depth is achieved by adjusting the volume space velocity and reaction temperature. The above reaction conditions are only exemplary. In order to control the conversion depth and isomerization depth in the first hydrogenation reactor during the reaction process so that the conversion rate of the fraction with an initial boiling point greater than 400°C is not greater than 40% and the isoparaffin content in the first reaction product is not less than 50% by weight, and in order to control the conversion depth in the second hydrogenation reactor during the reaction process so that the conversion rate of the fraction with an initial boiling point greater than 400°C is not less than 40%, those skilled in the art can make reasonable adjustments according to the specific reaction conditions. In addition, the reasonable operating range is also closely related to the specific performance of the catalyst used, and the above reaction conditions may vary according to changes in the specific catalyst used.

[0051] In some embodiments, in order to ensure long-term operation of the hydrogenation unit, a hydrogenation protection catalyst can be loaded on the top of the first hydrogenation reactor and the second hydrogenation reactor. Loading a hydrogenation protection catalyst to ensure long-term operation of the hydrogenation unit is well known to those skilled in the art. The type and loading amount of the hydrogenation protection catalyst are conventionally selected and will not be described in detail. For example, the hydrogenation protection catalyst includes, but is not limited to, FZC-100, FZC-101, FZC-102, FZC-102A, FZC-102B, FZC-103, FZC-103A, FZC-103B, FZC-105, FZC-106, and FZC-204 developed by Fushun Petrochemical Research Institute, and the loading amount is, for example, 10-35% of the volume of the hydroisomerization cracking catalyst in the corresponding hydrogenation reactor.

[0052] In some embodiments, a post-finishing agent may be loaded at the bottom of the first hydroisomerization cracking catalyst bed of the first hydrogenation reactor. This is a conventional operation in the art, and the type and loading amount of the post-finishing agent are conventionally selected. For example, the post-finishing agent includes one or more of 3962, FF-12, and FF-66 developed by Fushun Petrochemical Research Institute, and the loading amount is, for example, 10-30% of the volume of the first hydroisomerization cracking catalyst.

[0053] Unless otherwise specified or explained, the catalysts used in the first hydrogenation reactor and the second hydrogenation reactor of the present invention may also be selected from one or more catalysts in the prior art that can achieve the corresponding functions. They may be commercially available products or prepared by any existing method.

[0054] The "hydrogen-containing stream" introduced into the first and second hydrogenation reactors of the present invention refers to a stream capable of providing hydrogen, including any one or more of fresh hydrogen from a pipeline network, recycled hydrogen, hydrogen-rich gas, other gaseous streams capable of providing hydrogen, and other liquid streams capable of providing hydrogen. In the present invention, the "hydrogen-containing streams" introduced into the first and second hydrogenation reactors may be the same or different.

[0055] The present invention also provides a method for preparing white oil, comprising fractionating the first middle distillate and / or the second middle distillate obtained by the distillate hydrodearomatization method described above to obtain white oil. The distillate hydrodearomatization method described above can directly produce middle distillate with a low aromatic content, thereby easily and efficiently finely cutting the middle distillate to obtain light white oil, industrial white oil, cosmetic-grade white oil, and food-grade white oil with aromatic content that meets corresponding standards, with high yield. The fine-cutting operation specifically employed to finely cut the middle distillate to obtain white oil is a conventional method in the art, and refers to cutting through normal and vacuum towers to control the distillation temperature range to obtain narrow fractions. Narrow fractions with different distillation ranges are light white oils and industrial white oils of different grades, such as a distillation range of 160-195°C corresponding to W-40 light white oil.

[0056] Herein, the conversion rate of the fraction with an initial boiling point greater than 400°C = (content of the fraction with an initial boiling point greater than 400°C in the feedstock - content of the fraction with an initial boiling point greater than 400°C in the reaction product * reaction liquid yield) / content of the fraction with an initial boiling point greater than 400°C in the feedstock * 100%; in the formula, "feedstock" refers to the distillate oil used as the feedstock for the first hydrogenation reactor, and refers to the heavy oil with an initial boiling point greater than 400°C entering the second hydrogenation reactor;

[0057] Middle distillate oil yield = (content of fraction with a distillation range of 160-400° C. in the reaction product * reaction liquid yield - content of fraction with a distillation range of 160-400° C. in the raw material) / content of fraction with a distillation range of 160-400° C. in the raw material * 100%;

[0058] The reaction liquid yield refers to the mass ratio of the collected liquid product to the raw material.

[0059] The present invention is described below by way of examples. The raw oils used in the examples and comparative examples are all refined tail oils from Ningxia Coal Industry Coal-to-Liquids Company, and their properties are shown in Table 1.

[0060] Table 1

[0061]

[0062] Example 1

[0063] The distillate oil (i.e., feed oil) and a hydrogen-containing stream (hydrogen) are introduced into a first hydrogenation reactor, and the reaction effluent (i.e., the first reaction product) is sequentially separated and fractionated in a separation and fractionation unit to obtain naphtha, middle distillate oil with a distillation range of 160-400°C, and heavy oil with an initial boiling point greater than 400°C. The heavy oil fraction and the hydrogen-containing stream (hydrogen) are introduced into a second hydrogenation reactor, and the effluent from the second hydrogenation reactor (i.e., the second reaction product) is returned to the separation and fractionation unit for separation and fractionation.

[0064] The first hydroisomerization cracking catalyst used in the first hydrogenation reactor was prepared by the method disclosed in patent application CN110433819A, with reference to the steps and process conditions in Example 1 of CN110433819A. The pore volume of the amorphous silica-alumina used was 1.6 mL / g, and the specific surface area of ​​the amorphous silica-alumina was 500 m 2 / g, and the required amount of each raw material was weighed according to the needs of the catalyst to finally prepare the first hydroisomerization cracking catalyst used in this embodiment. Based on the total weight of the catalyst, the content of amorphous silica-alumina was 62 weight%, the content of aluminum oxide was 21 weight%, the content of tungsten oxide was 15.5 weight%, and the content of nickel oxide was 1.5 weight%.

[0065] The reaction conditions of the first hydrogenation reactor were controlled (hydrogen partial pressure of 6.4 MPa, reaction temperature of 320 °C, hydrogen-to-oil volume ratio of 800:1, volume space velocity of 1 h -1 ), so that the conversion rate of the fraction with an initial boiling point greater than 400° C. is equal to 38% and the isoparaffin content in the reaction effluent of the first hydrogenation reactor is 50% by weight.

[0066] The second hydroisomerization cracking catalyst used in the second hydrogenation reactor was FC14 (developed by Fushun Petrochemical Research Institute); the hydrorefining catalyst used in the second hydrogenation reactor was prepared by the equal volume impregnation method (the specific preparation process is described above and will not be repeated here), and its specific surface area is 220m 2 / g, pore volume is 0.51ml / g, infrared acid content is 0.83mmol / g, B acid / L acid ratio is 1.9; the chemical composition of the hydrorefining catalyst is 46.6wt% aluminum oxide, 8.1wt% nickel oxide, 12wt% molybdenum oxide, 30wt% tungsten oxide, 2wt% phosphorus pentoxide, and the remainder is impurities.

[0067] According to the flow direction of the reactant stream, the second hydroisomerization cracking catalyst and the hydrorefining catalyst were sequentially loaded into the second hydrogenation reactor, with the loading volume ratio of the two being 100:15.

[0068] By controlling the reaction conditions of the second hydrogenation reactor (hydrogen partial pressure of 6.4 MPa, reaction temperature of 345 ° C, hydrogen-to-oil volume ratio of 980:1, volume space velocity of 1.6 h -1 ), so that the conversion rate of the fraction with an initial distillation point greater than 400°C is 66%.

[0069] The obtained middle distillate oil with a distillation range of 160-400° C. and a yield of 82.3% was fractionated for further fine cutting to obtain light white oil and industrial white oil of different grades, as shown in Table 2.

[0070] The test methods specified in NB / SH / T0913-2015 "Light White Oil" and NB / SH / T0006-2017 "Industrial White Oil" were used to test various brands of white oil. The results are shown in Table 2.

[0071] Table 2

[0072]

[0073] Example 2

[0074] The middle distillate oil with a distillation range of 160-400° C. obtained in Example 1 was fractionated for further fine cutting to obtain food-grade white oil. The aromatic content of the obtained food-grade white oil meets the national food safety standard GB1886.216-2016. The results are shown in Table 3:

[0075]

[0076] Example 3

[0077] The middle distillate oil with a distillation range of 160-400° C. obtained in Example 1 was fractionated for further fine cutting to obtain cosmetic grade white oil. The aromatic content of the obtained cosmetic grade white oil complies with NB / SH / T0007-2015 "Cosmetic Grade White Oil". The results are shown in Table 4:

[0078] Table 4

[0079]

[0080]

[0081] Example 4

[0082] The raw oil is a mixture of cracked tail oil and refined tail oil from Ningxia Coal Industry Coal-to-Liquids Company. The mass ratio of cracked tail oil to refined tail oil is 1:1. The properties of refined tail oil are shown in Table 1, and the properties of cracked tail oil are shown in Table 5 below.

[0083] Table 5 Properties of distillate oil feedstock

[0084]

[0085] The distillate oil (i.e., feed oil) and a hydrogen-containing stream (hydrogen) are introduced into a first hydrogenation reactor, and the reaction effluent (i.e., the first reaction product) is sequentially separated and fractionated in a separation and fractionation unit to obtain naphtha, middle distillate oil with a distillation range of 160-400°C, and heavy oil with an initial boiling point greater than 400°C. The heavy oil fraction and the hydrogen-containing stream (hydrogen) are introduced into a second hydrogenation reactor, and the effluent from the second hydrogenation reactor (i.e., the second reaction product) is returned to the separation and fractionation unit for separation and fractionation.

[0086] The first hydroisomerization cracking catalyst used in the first hydrogenation reactor was prepared by the method disclosed in patent application CN110433819A, with reference to the steps and process conditions in Example 1 of CN110433819A. The amorphous silica-alumina used had a pore volume of 1.4 mL / g and a specific surface area of ​​370 m 2 / g, and the required amount of each raw material was weighed according to the needs of the catalyst to finally prepare the first hydroisomerization cracking catalyst used in this embodiment. Based on the total weight of the catalyst, the content of amorphous silica-alumina was 36 weight%, the content of aluminum oxide was 35.5 weight%, the content of tungsten oxide was 22.5 weight%, and the content of nickel oxide was 6 weight%.

[0087] By controlling the reaction conditions of the first hydrogenation reactor (hydrogen partial pressure of 6.4 MPa, reaction temperature of 319 °C, hydrogen-to-oil volume ratio of 800:1, volume space velocity of 1 h -1 ), so that the conversion rate of the fraction with an initial boiling point greater than 400° C. is equal to 40% and the isoparaffin content in the reaction effluent of the first hydrogenation reactor is 54.2% by weight.

[0088] The second hydroisomerization cracking catalyst used in the second hydrogenation reactor was FC16 (developed by Fushun Petrochemical Research Institute); the hydrorefining catalyst used in the second hydrogenation reactor was prepared by the equal volume impregnation method (the specific preparation process is described above and will not be repeated here), and its specific surface area is 372m 2 / g, pore volume is 1.1ml / g, infrared acid content is 0.34mmol / g, and the ratio of B acid / L acid is 1.5; the chemical composition of the hydrorefining catalyst is 12wt% nickel oxide, 16.6wt% molybdenum oxide, 28wt% tungsten oxide, and the balance is alumina support.

[0089] According to the flow direction of the reactant stream, the second hydroisomerization cracking catalyst and the hydrorefining catalyst were sequentially loaded into the second hydrogenation reactor, with the loading volume ratio of the two being 100:25.

[0090] By controlling the reaction conditions of the second hydrogenation reactor (hydrogen partial pressure of 6.4 MPa, reaction temperature of 345 ° C, hydrogen-to-oil volume ratio of 980:1, volume space velocity of 1.6 h -1 ), so that the conversion rate of the fraction with an initial distillation point greater than 400°C is 71%.

[0091] The obtained middle distillate oil with a distillation range of 160-400° C. and a yield of 76.6% was fractionated to further finely cut to obtain light white oil and industrial white oil of different grades, as shown in Table 6.

[0092] The test methods specified in NB / SH / T0913-2015 "Light White Oil" and NB / SH / T0006-2017 "Industrial White Oil" were used to test various brands of white oil. The results are shown in Table 6.

[0093] Table 6

[0094]

[0095]

[0096] Comparative Example 1

[0097] The process was carried out in accordance with Example 1, except that the reaction conditions of the first hydrogenation reactor and the second hydrogenation reactor were controlled so that the conversion rate of the fraction with an initial boiling point greater than 400° C. in the reaction effluent of the first hydrogenation reaction was 60%.

[0098] The obtained middle distillate oil with a distillation range of 160-400°C is fractionated to further finely cut to obtain light white oil and industrial white oil of different grades, see Table 7.

[0099] The test methods specified in NB / SH / T0913-2015 "Light White Oil" and NB / SH / T0006-2017 "Industrial White Oil" were used to test various brands of white oil. The results are shown in Table 7:

[0100] Table 7

[0101]

[0102] It can be found from Example 1 and Comparative Example 1 that changing the operation control method of the first hydrogenation reactor (conversion rate is 60%) leads to an increase in multiple cracking reactions in the first reaction zone, an increase in the content of produced aromatics, and a reduction in the yield of the middle fraction to 49.8%, resulting in a substantial increase in the aromatic content of the final product light white oil. Some brands of light white oil do not meet the NB / SH / T0913-2015 standard; 5# and 10# industrial white oils do not meet the technical requirements of industrial white oil (II) in NB / SH / T0006-2017, and the total yield of the white oil fraction is only 49.8%.

[0103] Comparative Example 2

[0104] Proceed with reference to Example 1, except that:

[0105] The first hydroisomerization cracking catalyst of the first hydrogenation reactor is loaded with a cracking agent containing molecular sieve (the preparation method of the cracking agent is the same as that of Example 1 in CN101450320B);

[0106] By controlling the reaction conditions of the first hydrogenation reactor, the conversion rate of the fraction with an initial boiling point greater than 400° C. is equal to 38%, but the isoparaffin content in the reaction effluent of the first hydrogenation reactor is 14.7% by weight.

[0107] By controlling the reaction conditions of the second hydrogenation reactor, the conversion rate of the fraction with an initial boiling point greater than 400° C. was set to 66%.

[0108] The obtained middle fraction with a distillation range of 160-400°C is fractionated for further fine cutting to obtain light white oil and industrial white oil of different grades, see Table 8.

[0109] The test methods specified in NB / SH / T0913-2015 "Light White Oil" and NB / SH / T0006-2017 "Industrial White Oil" were used to test various brands of white oil. The results are shown in Table 8.

[0110] Table 8

[0111]

[0112] In Comparative Example 2, a cracking agent containing molecular sieves was loaded into the first hydrogenation reactor. Although the conversion rate was the same, the isoparaffin content in the reaction effluent was 14.7% by weight, which also resulted in a significant increase in the aromatic content of the final product, light white oil. Some brands of light white oil did not meet the NB / SH / T0913-2015 standard, and 5# and 10# industrial white oils did not meet the technical requirements of industrial white oil (II) in NB / SH / T0006-2017, and the total yield of the middle distillate was only 45.8%.

[0113] It will be readily understood that the above embodiments are merely examples for clarity of description and are not intended to limit the present invention to these examples. Those skilled in the art will readily appreciate that other variations or modifications based on the above description are possible. It is not necessary and impossible to provide an exhaustive list of all possible implementations. Obvious variations or modifications arising therefrom remain within the scope of protection of the present invention.

Claims

1. A method for hydrodearomatization of distillate oil, characterized in that: The steps include: A distillate oil with a boiling range of 180-800° C. as a raw material is reacted with a hydrogen-containing stream in a first hydrogenation reactor in the presence of a first hydroisomerization catalyst to obtain a first reaction product, and the conversion depth and isomerization depth are controlled during the reaction so that the conversion rate of the fraction with an initial boiling point greater than 400° C. is no greater than 40% and the isoparaffin content in the first reaction product is no less than 50% by weight; the distillate oil as the raw material includes a mixture of one or more of synthetic light oil, synthetic wax full fraction, Fischer-Tropsch refined tail oil, and Fischer-Tropsch cracking tail oil from a Fischer-Tropsch synthesis process; Separating the first reaction product by fractional distillation to obtain a first middle distillate oil with a distillation range of 160 to 400° C. and a first heavy oil with an initial boiling point greater than 400° C. reacting the first heavy oil with a hydrogen-containing stream in a second hydrogenation reactor in the presence of a second hydroisomerization catalyst to obtain a second reaction product, and controlling the conversion depth during the reaction so that the conversion rate of the fraction having an initial boiling point greater than 400° C. is not less than 40%; Separating the second reaction product by fractional distillation to obtain a second middle distillate oil with a distillation range of 160 to 400° C. and a second heavy oil with an initial boiling point greater than 400° C. The first hydroisomerization cracking catalyst is a hydroisomerization cracking catalyst containing amorphous silica and aluminum; the first hydroisomerization cracking catalyst has a content of amorphous silica and aluminum of 32 to 64.5%, a content of aluminum oxide of 20 to 50%, a content of tungsten oxide of 5 to 30%, and a content of nickel oxide of 0.5 to 9%, based on the total weight of the catalyst; The second hydroisomerization cracking catalyst and the hydrorefining catalyst are sequentially loaded into the second hydrogenation reactor along the flow direction of the reaction materials; in the second hydrogenation reactor, the loading volume ratio of the second hydroisomerization cracking catalyst to the hydrorefining catalyst is 100:10-30; the second hydroisomerization cracking catalyst is a hydroisomerization cracking catalyst containing a molecular sieve; The reaction conditions of the first hydrogenation reactor include: reaction temperature 310-370 ° C, volume space velocity 1-4h -1 , hydrogen-oil volume ratio 500-1500:1, hydrogen partial pressure: 5-15MPa; the reaction conditions of the second hydrogenation reactor include: reaction temperature 320-375℃, volume space velocity 1-2h -1 , hydrogen-oil volume ratio 500-1500:1, hydrogen partial pressure: 5-15MPa.

2. The method according to claim 1, characterized in that The second heavy oil is circulated to the second hydrogenation reactor to participate in the reaction.

3. The method according to claim 1, characterized in that The first reaction product and the second reaction product are separated in separate devices, or the first reaction product and the second reaction product are separated in the same device.

4. The method according to claim 1, wherein The amorphous silica-alumina used in preparing the first hydroisomerization cracking catalyst has a pore volume of 1.0 to 1.6 mL / g and a specific surface area of ​​300 to 500 m 2 / g.

5. The method according to claim 1, wherein The first heavy oil and the hydrogen-containing stream are reacted in a second hydrogenation reactor in the presence of a second hydroisomerization cracking catalyst to obtain a second reaction product, and the conversion depth is controlled during the reaction so that the conversion rate of the fraction with an initial boiling point greater than 400° C. is not less than 65%.

6. The method according to any one of claims 1 to 5, characterized in that The hydrorefining catalyst includes a carrier and an active component supported on the carrier; based on the weight of the hydrorefining catalyst, the active component includes 10wt% to 50wt% of tungsten oxide, 5wt% to 30wt% of molybdenum oxide, and 0.5wt% to 20wt% of nickel oxide.

7. The method according to claim 6, characterized in that The carrier includes one or both of aluminum oxide and silicon oxide.

8. The method according to claim 6, characterized in that The specific surface area of ​​the hydrorefining catalyst is 100 to 400 m 2 / g, the pore volume is 0.3~1.2ml / g, the infrared acid content is 0.3~1.0mmol / g, and the ratio of B acid / L acid is 1.5~2.

0.

9. A method for preparing white oil, characterized in that: The white oil is obtained by fractionating the first middle distillate oil and / or the second middle distillate oil obtained by the method according to any one of claims 1 to 8.

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