CN116408090B

CN116408090B - Hydrogenation catalyst, preparation method thereof, method and system for producing solvent oil by reforming raffinate oil

Info

Publication number: CN116408090B
Application number: CN202111674373.3A
Authority: CN
Inventors: 鞠雅娜; 张然; 张雅琳; 袁晓亮; 吕忠武; 宋绍彤; 钟海军; 侯远东; 雷俊伟; 王嘉祎
Original assignee: Petrochina Co Ltd
Current assignee: Petrochina Co Ltd
Priority date: 2021-12-31
Filing date: 2021-12-31
Publication date: 2024-10-15
Anticipated expiration: 2041-12-31
Also published as: KR20240129045A; CN116408090A; WO2023125044A1

Abstract

The invention provides a hydrogenation catalyst, a preparation method thereof, a method and a system for producing solvent oil by reforming raffinate oil. The hydrogenation catalyst comprises a carrier, a first active metal component and a second active metal component, wherein the first active metal component comprises nickel, and the weight content of the first active metal component in the catalyst is 40-70%. The invention also provides a preparation method of the catalyst. The invention provides a method for producing solvent oil by reforming raffinate oil, which utilizes the catalyst to carry out hydrogenation benzene removal and olefin removal reaction on reformed raffinate oil after desulfurization and dehydration. The invention also provides a system for producing solvent oil by reforming raffinate oil, which can realize the method. The method is suitable for separating and utilizing the reformed raffinate oil with higher benzene and olefin content, can utilize the reformed raffinate oil to produce various solvent oils, and effectively improves the added value of the reformed raffinate oil.

Description

Hydrogenation catalyst, preparation method thereof, method and system for producing solvent oil by reforming raffinate oil

Technical Field

The invention relates to the technical field of reforming raffinate oil treatment, in particular to a hydrogenation catalyst, a preparation method thereof, a method and a system for producing solvent oil by reforming raffinate oil.

Background

The reformed raffinate oil is a byproduct in aromatic hydrocarbon production, is a good petrochemical raw material because of low content of sulfur, nitrogen, heavy metals and other impurities, can separate different high-added-value components according to different boiling points, and is suitable for producing high-quality solvent oil, high-added-value hexane oil, isohexane oil and the like.

The reformed raffinate oil contains a small amount of unsaturated components such as olefin, aromatic hydrocarbon and the like, and the property of the raffinate oil serving as the solvent oil with high added value is seriously affected. Such as: n-hexane is mainly used as a solvent for olefin polymerization such as propylene, an edible vegetable oil extracting agent, a rubber and paint solvent, a pigment diluent and the like, and has definite limitation on unsaturated hydrocarbons such as benzene and the like; the No. 6 solvent oil is an organic solvent for producing edible oil, part of the edible oil remains in the production process, and the carcinogenic aromatic hydrocarbon in the solvent oil remains in the edible oil, so that the organic solvent is harmful to human bodies, and the aromatic hydrocarbon content in the No. 6 solvent oil must be strictly controlled; the pentane foaming agent is mainly used as a solvent, for manufacturing artificial ice, anesthetic and the like, and the bromine index is definitely required to be less than or equal to 100mgBr/100g. Therefore, in order to reasonably utilize the reformed raffinate oil and endow the reformed raffinate oil with higher economic value, the process of hydrodeolefination and dearomatization must be carried out to remove unsaturated hydrocarbons in the reformed raffinate oil. At present, representative hydrodeolefine and dearomatization catalysts are mainly divided into two major categories of noble metal and non-noble metal, and three metal catalysts of Ni, pt and Pd are more applied. Pd hydrogenation performance K _Pd is 1/7 of metal Ni, pt activity K _Pt is 2.5 times of Ni, but the price of palladium and platinum is hundreds of times higher than that of nickel, and the requirements on sulfur, arsenic and other toxic substances in raw materials are higher,

At present, the reforming raffinate oil is used for producing high added value products, and the adopted technological processes are different according to the raw material properties and different product requirements, and mainly comprise two steps of hydrogenation, fractionation and hydrogenation. When the benzene and olefin contents are higher, a process flow of hydrogenation and fractionation is basically adopted to ensure that the benzene content of the product meets the standard requirement; when the benzene and olefin contents are low, a process flow of fractionation and hydrogenation is basically adopted, the hydrogenation operation load is smaller, but the requirement on the purity of hydrogen is higher. The number of tower plates and the number of towers are different due to the existence of component azeotropic phenomenon in the fractionation process. Meanwhile, nickel hydrogenation dealkenation and debenzolization catalysts are poor in thermal stability at high temperature, niO particles are easy to agglomerate and sinter, the problems of poor dispersity, poor utilization rate and the like exist, and impurities such as sulfur, nitrogen, heavy metals and the like in raw materials easily cause catalyst poisoning, influence the service life of the catalyst, increase the operation cost and require pretreatment of the raw materials. Therefore, the existing process flow often has the problems of low utilization rate of the reformed raffinate oil, single processed product, low product yield, short operation period of the hydrogenation catalyst, insensitive temperature rising and the like, and needs to comprehensively consider various factors such as raw material properties, product types, catalyst stability, sulfur resistance, fractionation precision and the like to realize the efficient utilization of the reformed raffinate oil.

For the process of fractionation and hydrogenation, the components with the carbon number of C5 and below are not subjected to hydrogenation saturation treatment, so that benzene and bromine index are easy to be disqualified, the requirement of a pentane foaming agent can not be met, and the process is not suitable for raw materials with higher benzene content; the material flow II is a C6 mixture, and has the problem of low isomerization conversion rate; the light hydrocarbon component in the isohexane is higher in the subsequent fractionation process because the hydrogenation and isomerization products are not subjected to gas-liquid separation; the raw materials entering the hydrogenation reactor are not subjected to desulfurization treatment, which is easy to cause poisoning and deactivation of the hydrogenation catalyst.

For the process of hydrogenation before separation, when the rectification treatment is carried out, sulfolane, water and impurities are discharged from the bottom of the tower, so that the loss of part of heavy fractions is easy to influence the yield of the No. 120 solvent oil, and meanwhile, the rectification separation of the C6 product at the top of the rectification tower is not carried out, so that the yield of normal hexane is easy to be lower, high-purity normal hexane with the price superior to that of the No. 6 solvent oil cannot be produced, and the high-efficiency utilization of raffinate oil is not realized.

Meanwhile, the process of producing solvent oil by reforming raffinate oil is reported to adopt two sections of hydrogenation process flows for hydrogenation and benzene removal, wherein the first section adopts a hydrodesulfurization catalyst to remove sulfolane so as to protect a subsequent catalyst, and the second section adopts a high-nickel hydrogenation catalyst to remove benzene, but because the hydrodesulfurization catalyst needs to be pre-vulcanized first, the second section of catalyst needs to be pre-reduced first, the two sections of catalyst cannot be operated simultaneously, the starting process is complicated, and residual vulcanized oil in the first section also has an influence on the second section of catalyst, so that a large amount of raw oil is required to be washed, and the starting period of the device is influenced.

Disclosure of Invention

In order to solve the problems, the invention aims to provide a hydrogenation catalyst, a preparation method thereof, a method for producing solvent oil by reforming raffinate oil and a system thereof. The method is suitable for separating and utilizing the reformed raffinate oil with higher benzene and olefin content, can utilize the reformed raffinate oil to produce various solvent oils, and effectively improves the added value of the reformed raffinate oil.

In order to achieve the above object, the present invention provides a hydrogenation catalyst comprising a support, and a first active metal component and a second active metal component supported on the support, wherein the metal element of the first active metal component comprises nickel, the weight content of the first active metal component in the catalyst is 40-70% based on the total weight of the catalyst, the weight content of the support (based on oxide) in the catalyst is 20-59%, and the balance is the second active metal component.

In the specific embodiment of the invention, compared with the existing nickel-based catalyst and hydrogenation catalyst, the hydrogenation catalyst provided by the invention has higher nickel content, higher dispersity of Ni metal which can reach more than 10%, and more hydrogenation active centers, so that the hydrogenation catalyst has better benzene and olefin removal effect and more thorough removal.

In the above hydrogenation catalysts, the weight of the first active metal component in the catalyst is generally controlled to be in the range of 40-70%, such as 40%, 45%, 50%, 55%, 60%, 65%, 70% or any two of these.

In the above hydrogenation catalyst, the weight content of the carrier in the catalyst is in the range of 20 to 59%, for example 20%, 25%, 30%, 35%, 40%, 45%, 50%, 55%, 59% or any two thereof.

In particular embodiments of the invention, the specific surface area of the hydrogenation catalyst is generally in the range of 200-500m ²/g, such as 200m²/g、250m²/g、300m²/g、350m²/g、400m²/g、450m²/g、500m²/g or any two of these.

In particular embodiments of the invention, the pore volume of the hydrogenation catalyst is generally in the range of 0.3 to 0.6cm ³/g, e.g., 0.3cm³/g、0.35cm³/g、0.4cm³/g、0.45cm³/g、0.5cm³/g、0.55cm³/g、0.6cm³/g or any two of these.

In a specific embodiment of the present invention, the carrier is generally an oxide, and may include, for example, one of alumina, silica, titania, and ceria, or may include a mixture of alumina and at least one of silica, titania, and ceria. That is, the support may include one or a combination of two or more of alumina, silica, titania, and ceria, and the content of alumina is not 0.

In particular embodiments of the present invention, the first and second active metal components of the above-described catalysts are typically present as metal oxides.

In a specific embodiment of the present invention, the metal element of the second active metal component may include one or a combination of two or more of copper, lanthanum, magnesium, and the like. In a specific embodiment of the present invention, the nickel content, nickel dispersity and hydrogenation active center content of the catalyst can be improved by controlling the coprecipitation process and precipitation pH value of the metal active components during the preparation process of the catalyst.

The invention further provides a preparation method of the hydrogenation catalyst, which comprises the following steps:

the temperature of the slurry of the carrier precursor is enabled to reach the precipitation reaction temperature, and the pH value of the slurry of the carrier precursor is adjusted to reach the alkaline pH value (the adjustment of the pH value can be realized by adding a precipitator), so as to obtain a first mixed system;

Simultaneously adding a first active metal component precursor, a second active metal component precursor and a precipitator into the first mixed system, and keeping the pH value of the system at an alkaline pH value to obtain a second mixed system;

And (3) aging the second mixed system, and roasting the aged product to obtain the hydrogenation catalyst.

In the above preparation method, the carrier precursor is converted into the carrier of the hydrogenation catalyst after the drying and baking. In particular embodiments, the support precursor may comprise pseudoboehmite, which is then dried and calcined to convert to alumina.

In the preparation method, the nickel content, the nickel dispersity and the hydrogenation active center content in the catalyst can be improved by adjusting the precursor type of the first active metal component, the precursor type of the second active metal component and the precursor type of the precipitant.

In the above preparation method, the precipitant may be used not only to adjust the pH value, but also to precipitate the metal element in the first active metal component precursor and the metal element in the second active metal component precursor. In particular, the precipitant may comprise an alkaline precipitant, preferably comprising a soluble carbonate, for example comprising potassium carbonate and/or sodium carbonate.

In the above preparation method, the amount of the precipitant may be determined according to the adjustment of pH and the amount required to precipitate the metal element in the first active metal component precursor and the metal element in the second active metal component precursor.

In the above-described production method, the first active metal component precursor may be a compound of the metal element in the first active metal component, for example, a metal salt including the metal element in the first active metal component, preferably a soluble metal salt including the metal element in the first active metal component, for example, a nitrate including the metal element in the first active metal component, such as nickel nitrate, or the like.

In the above-described production method, the second active metal component precursor may be a compound of the metal element in the second active metal component, for example, a metal salt including the metal element in the second active metal component, preferably a soluble metal salt including the metal element in the second active metal component, for example, a nitrate including the metal element in the second active metal component, such as magnesium nitrate, or the like.

In the preparation method, the amount of the carrier precursor, the first active metal component precursor and the second active metal component precursor corresponds to the weight content of the carrier, the first active metal component and the second active metal component in the hydrogenation catalyst.

In a specific embodiment of the invention, the first active metal component precursor and the second active metal component precursor may be added to the first mixed system as a mixed solution of the two; the precipitant is also typically added to the first mixed reaction system in solution. By the parallel flow adding mode of the mixed solution of the first active metal component precursor and the second active metal component precursor and the precipitant solution and controlling the dripping rate of the two mixed solutions, the nickel content, the nickel dispersity and the hydrogenation active center content in the catalyst can be effectively controlled. In a specific embodiment, the dropping rate of the mixed solution of the first active metal component precursor and the second active metal component precursor is generally controlled to be 20 to 100ml/min, and the dropping rate of the precipitant solution is generally controlled to be 10 to 300ml/min.

In the above preparation method, the alkaline pH is in the range of 8 to 11, for example 8, 9, 10, 11 or any two thereof.

In the above preparation method, the precipitation reaction temperature is in the range of 40 to 90 ℃, for example 40 ℃, 50 ℃, 60 ℃, 70 ℃, 80 ℃, 90 ℃ or any two of them.

In the above preparation method, the temperature of the aging treatment and the precipitation reaction temperature may be the same or different. The temperature of the aging treatment is generally in the range of 40-90 ℃, for example 40 ℃, 50 ℃, 60 ℃, 70 ℃, 80 ℃, 90 ℃ or any two thereof.

In the above preparation method, the time of the aging treatment is generally controlled to be in the range of 1 to 5 hours, for example, 1 hour, 2 hours, 3 hours, 4 hours, 5 hours or any two of them.

In the above preparation method, the temperature of the calcination may be generally controlled to be in the range of 300 to 600 ℃, for example 300 ℃, 350 ℃, 400 ℃, 450 ℃, 500 ℃, 550 ℃, 600 ℃, or any two of them.

In the above preparation method, the time of the calcination is generally in the range of 1 to 5 hours, for example, 1 hour, 2 hours, 3 hours, 4 hours, 5 hours or any two of them.

In a specific embodiment of the present invention, the preparation method generally further comprises the operations of filtering, washing and drying sequentially after the aging treatment and before the calcination.

In the above preparation method, the washing is generally stopped when the filtrate reaches neutrality. The filter cake obtained is subjected to subsequent drying and calcination.

In the above preparation method, the temperature of the drying may be in the range of 100 to 130 ℃, for example, 100 ℃, 110 ℃, 120 ℃, 130 ℃, or any two thereof.

In the above preparation method, the drying time is generally in the range of 1 to 5 hours, for example, 1 hour, 2 hours, 3 hours, 4 hours, 5 hours or any two thereof.

According to a specific embodiment of the present invention, the preparation method of the hydrogenation catalyst may include:

Heating slurry formed by a carrier precursor and water to a precipitation reaction temperature of 40-90 ℃ while stirring, and adding a precipitant aqueous solution into the slurry to enable the pH value of the system to be 8-11, so as to obtain a first mixed system; simultaneously adding a mixed aqueous solution of a first active metal component precursor and a second active metal component precursor (the mixed aqueous solution can be formed by mixing an aqueous solution of the first active metal component precursor and an aqueous solution of the second active metal component precursor) and an aqueous solution of a precipitant into a first mixed system while stirring, and keeping the pH value of the system to be 8-11 to obtain a second mixed system; aging the second mixed system at 40-90 ℃ for 1-5h, filtering, washing until the filtrate is neutral, drying the filter cake, and roasting at 300-600 ℃ for 1-5h to obtain the hydrogenation catalyst.

The invention also provides a method for producing solvent oil by reforming raffinate oil, which comprises the following steps:

s2, catalyzing the reformed raffinate oil subjected to desulfurization and dehydration by using a hydrogenation catalyst to carry out hydrogenation and debenzolization and olefin removal reactions to obtain a reaction product; wherein the hydrogenation catalyst comprises the hydrogenation catalyst;

s3, carrying out gas-liquid separation on the reaction product obtained in the step S2;

S4, rectifying the separated liquid to respectively distillate a C5 component, a C6 component and more than C7 components;

S5, rectifying the C6 component to respectively distillate cyclohexane and a mixture of normal hexane and isohexane;

s6, rectifying the mixture of the normal hexane and the isohexane, and respectively distilling off the isohexane and the normal hexane to finish the production of the solvent oil.

In a specific embodiment of the invention, the content of sulfur in the non-desulfurized and dehydrated reformed raffinate oil is below 20mg/kg, the content of benzene is 3-5v%, the content of olefin is 1-3v%, and the total content of normal hexane and isohexane is more than 50v%.

In a specific embodiment of the present invention, the above method further comprises: s1, desulfurizing and dehydrating the heavy raffinate oil, wherein the S content of the product is less than or equal to 1mg/kg.

In the specific embodiment of the invention, in S1, the space velocity in the desulfurization and dehydration process is 1-10h ^-1;

In particular embodiments of the present invention, the adsorbent employed in the desulfurization and dehydration process in S1 comprises a molecular sieve adsorbent, such as one or a combination of two or more of a 3A molecular sieve, a 4A molecular sieve, a 5A molecular sieve, and a 13A molecular sieve;

in particular embodiments of the present invention, the molecular sieve adsorbent used in S1 generally has a specific surface area of 400-700m ²/g. The pore volume of the molecular sieve adsorbent is generally 0.05-0.30cm ³/g.

In a specific embodiment of the present invention, in S2, the hydrogenation catalyst may be effective to remove benzene and olefins from the reformate residue. According to a specific embodiment of the present invention, in the reaction product obtained in S2, the benzene content is not more than 0.01% and the bromine index is not more than 50mgBr/100g.

In the specific embodiment of the present invention, in S2, the conditions of the hydrodebenzene and dealkenation reaction may be generally controlled as follows: the temperature is 90-200 ℃, the reaction pressure is 1.0-3.0MPa, and the hydrogen-oil volume ratio is 100-300: 1. volume airspeed 1.0-3.0h ^-1;

In a specific embodiment of the invention, in S4, the pressure of the rectification is generally between 0 and 0.2MPa.

In a specific embodiment of the invention, the rectification process of S4 can be completed in a rectification column, and the temperature at the top of the rectification column can be 30-50 ℃, the temperature at the bottom of the rectification column can be 90-110 ℃, and the pressure in the column can be 0-0.2MPa.

In a specific embodiment of the invention, in S5, the pressure of the rectification is generally from 0 to 0.1MPa.

In a specific embodiment of the invention, the rectification process of S5 can be completed in a rectification column, the top temperature of the rectification column is 55-75 ℃, the bottom temperature of the rectification column is 70-90 ℃, and the pressure in the column is 0-0.1MPa.

In a specific embodiment of the invention, in S6, the pressure of the rectification is 0-0.1MPa.

In a specific embodiment of the invention, the rectification process of S6 can be completed in a rectification column, wherein the temperature of the top of the rectification column is 50-70 ℃, the temperature of the bottom of the rectification column is 60-80 ℃, and the pressure in the column is 0-0.1MPa.

In the specific embodiment of the invention, the isohexane obtained in the step S6 can be directly output as No. 6 solvent oil. According to actual needs, the isohexane obtained in S6 can be converted into n-hexane through a normal structuring reaction, and the purity of the n-hexane product is more than or equal to 80%.

In a specific embodiment of the invention, the method further comprises: s7, converting the isohexane obtained in the step S6 into normal hexane through a normal structuring reaction.

In a specific embodiment of the present invention, in S7, the conditions of the orthosteric reaction are: the reaction temperature is 200-400 ℃, the reaction pressure is 1.0-3.0MPa, and the hydrogen-oil volume ratio is 100-300: 1. the volume space velocity is 1.0-3.0h ^-1.

In a specific embodiment of the present invention, in S7, the catalyst of the orthosteric reaction may be a molecular sieve catalyst. In particular, the molecular sieve catalyst may include a third active metal component, a non-molecular sieve support, and a molecular sieve support.

In the molecular sieve catalyst, the molecular sieve carrier comprises one or more than two of MOR, MCM-41, ZSM-22 and SAPO-11.

In the above molecular sieve catalyst, the metal element of the third active metal component may include an element of group VIB and/or VIII. In particular embodiments, the third active metal component is typically present in the form of a metal oxide.

In the above molecular sieve catalyst, the non-molecular sieve support comprises alumina.

In the above molecular sieve catalyst, the weight of the molecular sieve carrier is 10 to 80% of the total weight of the catalyst, the weight of the third active metal component (calculated as oxide) is 0.01 to 5% (e.g., 0.1 to 0.5%) of the total weight of the catalyst, based on 100% of the total weight of the molecular sieve catalyst, and the balance is the non-molecular sieve carrier.

In a specific embodiment of the present invention, in S7, the purity of isohexane as a raw material for the orthosteric reaction is 90% or more, preferably 95% or more.

In a specific embodiment of the present invention, the method for producing solvent oil by reforming raffinate oil may specifically comprise:

S1, desulfurizing and dehydrating the reformed raffinate oil by using a molecular sieve adsorbent;

s3, performing gas-liquid separation on the reaction product obtained in the step S2, and outputting hydrogen-rich gas and liquid;

s4, rectifying the separated liquid to respectively distillate a C5 component (which can be used as a pentane foaming agent), a C6 component and a component above C7 (which can be used as 120# solvent oil);

S6, rectifying the mixture of the normal hexane and the isohexane to respectively distillate the isohexane and the normal hexane;

s7, converting the isohexane obtained in the S6 into normal hexane through normal structuring reaction, and completing the production of solvent oil.

The invention also provides a system for producing the solvent oil by reforming the raffinate oil, and the system can realize the method for producing the solvent oil by reforming the raffinate oil.

In a specific embodiment of the invention, the system for producing solvent oil by reforming raffinate oil comprises a desorption desulfurization reactor, a hydrogenation reactor, a first gas-liquid separator, a first rectifying tower, a second rectifying tower and a third rectifying tower which are connected in sequence.

The connection relation of each device in the system can be as follows: the inlet of the desorption desulfurization reactor is used for receiving the reformed raffinate oil, and is generally positioned at the top of the tower;

the outlet of the desorption desulfurization reactor positioned at the bottom of the tower is connected with the inlet (positioned at the top of the tower) of the hydrogenation reactor;

The outlet of the hydrogenation reactor positioned at the bottom of the tower is connected with the inlet (positioned at the side line) of the first gas-liquid separator;

The outlet of the first gas-liquid separator positioned at the bottom of the tower is connected with the inlet (positioned at the side line) of the first rectifying tower;

the outlet of the first rectifying tower positioned on the measuring line is connected with the inlet of the second rectifying tower;

the outlet of the second rectifying tower positioned at the top of the tower is connected with the inlet of the third rectifying tower;

and the outlet of the third rectifying tower positioned at the top of the tower is used for outputting isohexane.

In the system, the outlet of the first gas-liquid separator positioned at the top of the tower is used for outputting hydrogen, the outlet of the first rectifying tower positioned at the top of the tower is used for outputting components with more than C7, the outlet of the first rectifying tower positioned at the bottom of the tower is used for outputting components with more than C7, the outlet of the first rectifying tower positioned at the line is used for outputting components with more than C6, the outlet of the second rectifying tower positioned at the top of the tower is used for outputting a mixture of normal hexane and isohexane, the outlet of the second rectifying tower positioned at the bottom of the tower is used for outputting cyclohexane, the outlet of the third rectifying tower positioned at the bottom of the tower is used for outputting normal hexane, and the outlet of the third rectifying tower positioned at the top of the tower is used for outputting isohexane.

In a specific embodiment of the invention, the system for producing solvent oil by reforming raffinate oil further comprises a orthographic formation reactor and a second gas-liquid separator, wherein the inlet of the orthographic formation reactor is connected with the outlet of the third rectifying tower at the top of the tower, the outlet of the orthographic formation reactor is connected with the inlet (at a side line) of the second gas-liquid separator, the outlet of the second gas-liquid separator at the top of the tower is used for outputting hydrogen, and the outlet of the second gas-liquid separator at the bottom of the tower is used for outputting n-hexane.

In particular embodiments of the present invention, the above-described system may further comprise a hydrogen line for replenishing hydrogen to the hydrogenation reactor and/or the orthographic reactor. The hydrogen line may be provided with a recycle hydrogen compressor, the inlet of which may be connected to the gas outlet of the first gas-liquid separator and/or the second gas-liquid separator, i.e. the compressor may be fed by the first gas-liquid separator, the second gas-liquid separator.

The invention has the beneficial effects that:

1. According to the method provided by the invention, the raw materials are pretreated by filling the molecular sieve adsorbent in the adsorption reactor, and desulfurization, dehydration and impurity removal are carried out, so that the problem of poisoning of a subsequent hydrogenation and benzene removal catalyst caused by sulfolane in the reforming raffinate oil raw materials in the prior art is solved, the service life of the subsequent hydrogenation and olefin removal high-nickel hydrogenation catalyst of a benzene removal unit is effectively protected, the running period of the device is prolonged, the running cost increase caused by frequent replacement of the catalyst is reduced, and the separated components above C7 meet the sulfur content requirement of the 120# solvent oil; meanwhile, the problems that the device starting process is complicated and the residual vulcanized oil is difficult to replace and the device starting period is influenced because the hydrodesulfurization catalyst needs to be presulfided in the prior art are avoided.

2. According to the method provided by the invention, the rectifying units can be arranged step by step from the molecular oil refining angle according to the different compositions and boiling points of raw materials, the normal structuring reaction of isohexane is utilized, the purity and the yield of normal hexane are improved, the normal hexane, pentane foaming agent and 120# solvent oil are produced, and the high-efficiency utilization of reformed raffinate oil is realized. The problems of low utilization rate of reformed raffinate oil, single processed product, low product yield and the like are effectively solved.

3. The nickel-based hydrogenation catalyst provided by the invention can realize high-content nickel active component dispersion on the carrier alumina, and the hydrogenation olefin removal and benzene removal reaction of the heavy raffinate oil is carried out, so that the operation period of a hydrogenation device is ensured.

Drawings

FIG. 1 is a schematic diagram of a process flow diagram and system connection for producing solvent oil from reformate according to the present invention.

Symbol description

The device comprises an adsorption desulfurization reactor 1, a hydrogenation reactor 2, a first gas-liquid separator 3, a first rectifying tower 4, a second rectifying tower 5, a third rectifying tower 6, a orthosteric reactor 7, a second gas-liquid separator 8 and a circulating hydrogen compressor 9.

Detailed Description

The technical solution of the present invention will be described in detail below for a clearer understanding of technical features, objects and advantageous effects of the present invention, but should not be construed as limiting the scope of the present invention.

In the following examples and comparative examples, the metal dispersity of the hydrogenation catalyst was measured by a static chemisorber using H ₂ as the adsorption gas, and the specific measurement method was as follows:

degassing the sample at 130 ℃, reducing the sample by H ₂ at 400 ℃, cooling to 35 ℃, performing first saturated adsorption, vacuumizing, and performing second saturated adsorption. The difference between the saturated adsorbtions is the volume of chemisorbed hydrogen.

The metal dispersity calculation formula:

in the formula, V _H2 is the volume of chemisorbed hydrogen, the unit is L, W is the total mass of the catalyst, the unit is g, P is the percentage of Ni, and M _Ni is the molecular mass of Ni.

Example 1a

The embodiment provides a hydrogenation catalyst, and the preparation method thereof comprises the following steps:

Step 1, mixing 11.50g of pseudo-boehmite powder (containing 70wt% of Al ₂O₃) with 40ml of deionized water and uniformly stirring to obtain first slurry;

58.16g of nickel nitrate hexahydrate is dissolved in deionized water, and the volume is fixed to 200ml to prepare 1mol/L nickel nitrate solution; 12.82g of magnesium nitrate hexahydrate was dissolved in deionized water to a volume of 50ml to prepare a 1mol/L magnesium nitrate solution. Weighing 184.7ml of 1mol/L nickel nitrate solution and 28.5ml of 1mol/L magnesium nitrate solution, and uniformly mixing to obtain a second solution;

207.32g of anhydrous potassium carbonate is dissolved in deionized water, the volume is fixed to 1000ml, 1.5mol/L potassium carbonate solution is prepared, and 665.3ml of 1.5mol/L potassium carbonate solution is measured as a third solution.

Step2, adding the first slurry into a reaction kettle, starting stirring speed of 200rmp, heating to a reaction temperature of 60 ℃, adding 1.3ml of a third solution into the reaction kettle, and regulating the pH of the first slurry to 10 to form a first mixed system;

And under the stirring condition, respectively adding the second solution and the third solution into the first mixed system in parallel flow at 35ml/min and 109.2ml/min, controlling the pH value to be constant at 10, and obtaining a second mixed system after the second solution and the third solution are added dropwise.

And 3, stirring and aging the second mixed system at 60 ℃ for 2 hours, filtering to obtain a filter cake after aging, washing to neutrality, drying the filter cake in a baking oven at 120 ℃ for 4 hours, grinding, and roasting the filter cake in a muffle furnace at 500 ℃ for 5 hours to obtain the nickel hydrogenation catalyst.

In the hydrogenation catalyst, the mass content of the nickel oxide is 60%, and the mass content of the magnesium oxide is 5%.

The catalyst of this example was tested to give a Ni metal dispersion of 15.08%, a specific surface area of 251m ²/g, and a pore volume of the hydrogenation catalyst of 0.41cm ³/g.

Example 1b

Step 1, mixing and uniformly stirring 16.10g of pseudo-boehmite powder (containing 70wt% of Al ₂O₃), 2.30g of silicon oxide and 55ml of deionized water to obtain a first slurry;

72.70g of nickel nitrate hexahydrate is dissolved in deionized water, and the volume is fixed to 500ml to prepare 0.5mol/L nickel nitrate solution; 4.33g of lanthanum nitrate hexahydrate is dissolved in deionized water to a volume of 20ml to prepare a 0.5mol/L lanthanum nitrate solution. Weighing 246.3ml of 0.5mol/L nickel nitrate solution and 2.8ml of 0.5mol/L lanthanum nitrate solution, and uniformly mixing to obtain a second solution;

15.81g of ammonium bicarbonate was dissolved in deionized water to a volume of 200ml to prepare a 1mol/L ammonium bicarbonate solution. And measuring 191.1ml of 1mol/L ammonium bicarbonate solution as a third solution.

Step 2, adding the first slurry into a reaction kettle, starting stirring speed of 200rmp, heating to a reaction temperature of 80 ℃, adding 1.2ml of a third solution into the reaction kettle, and regulating pH of the first slurry to 8 to form a first mixed system;

And under the stirring condition, respectively adding the second solution and the third solution into the first mixed system in parallel flow at 20ml/min and 15ml/min, controlling the pH value to be constant at 8, and obtaining a second mixed system after the second solution and the third solution are added dropwise.

And 3, stirring and aging the second mixed system at 80 ℃ for 3 hours, filtering to obtain a filter cake after aging, washing to neutrality, drying the filter cake in a baking oven at 120 ℃ for 4 hours, grinding, and roasting the filter cake in a muffle furnace at 500 ℃ for 5 hours to obtain the nickel hydrogenation catalyst.

In the hydrogenation catalyst, the mass content of the oxide of nickel is 70%, and the mass content of the oxide of lanthanum is 1%.

The catalyst of this example was tested to give a Ni metal dispersion of 16.18%, a specific surface area of 246m ²/g, and a pore volume of the hydrogenation catalyst of 0.40cm ³/g.

Example 1c

58.16g of nickel nitrate hexahydrate is dissolved in deionized water, and the volume is fixed to 200ml to prepare 1mol/L nickel nitrate solution; dissolving 4.83g of copper nitrate trihydrate in deionized water, and fixing the volume to 20ml to prepare 1mol/L copper nitrate solution; weighing 184.7ml of 1mol/L nickel nitrate solution and 14.5ml of 1mol/L copper nitrate solution, and uniformly mixing to obtain a second solution;

207.32g of anhydrous potassium carbonate is dissolved in deionized water, the volume is fixed to 1000ml, 1.5mol/L potassium carbonate solution is prepared, and 621.4ml of 1.5mol/L potassium carbonate solution is measured as a third solution.

Step2, adding the first slurry into a reaction kettle, starting stirring speed of 100rmp, heating to a reaction temperature of 60 ℃, adding 1.3ml of a third solution into the reaction kettle, and regulating the pH of the first slurry to 10 to form a first mixed system;

And under the stirring condition, respectively adding the second solution and the third solution into the first mixed system in parallel flow at 35ml/min and 109ml/min, controlling the pH value to be constant at 10, and obtaining a second mixed system after the second solution and the third solution are added dropwise.

And 3, stirring and aging the second mixed system at 60 ℃ for 2 hours, filtering to obtain a filter cake after aging, washing to neutrality, drying the filter cake in a baking oven at 120 ℃ for 4 hours, grinding, roasting in a muffle furnace at 500 ℃ for 5 hours, and reducing at 400 ℃ for 10 hours under high-purity H ₂ to obtain the nickel hydrogenation catalyst.

In the hydrogenation catalyst, the mass content of the nickel oxide was 60%, and the mass content of the copper oxide was 5%.

The catalyst of this example was tested to give a Ni metal dispersion of 14.74%, a specific surface area of 248m ²/g, and a pore volume of the hydrogenation catalyst of 0.42cm ³/g.

The process for producing solvent oil by reforming raffinate oil in examples 2a to 2e is carried out in a system shown in fig. 1, which comprises a desorption desulfurization reactor 1, a hydrogenation reactor 2, a first gas-liquid separator 3, a first rectifying column 4, a second rectifying column 5, a third rectifying column 6, a normal-construction reactor 7, a second gas-liquid separator 8, and a recycle hydrogen compressor 9.

Wherein, the top of the desorption desulfurization reactor 1 is provided with an inlet, and the bottom of the desorption desulfurization reactor is provided with an outlet; an inlet is arranged at the top of the hydrogenation reactor 2 an outlet is arranged at the bottom of the tower; the measuring line of the first gas-liquid separator 3 is provided with an inlet, and the bottom and the top of the tower are respectively provided with an outlet; the measuring line of the first rectifying tower 4 is provided with an inlet, and the tower top and the tower bottom are respectively provided with an outlet; the measuring line of the second rectifying tower 5 is provided with an inlet, and the tower top and the tower bottom are respectively provided with an outlet; the measuring line of the third rectifying tower 6 is provided with an inlet, and the top and the bottom of the tower are respectively provided with an outlet; the top of the orthosteric reactor 7 is provided with an inlet, and the bottom of the orthosteric reactor is provided with an outlet; the measuring line of the second gas-liquid separator 8 is provided with an inlet, and the top and the bottom of the tower are respectively provided with an outlet. The recycle hydrogen compressor 9 is provided with an inlet and an outlet.

Specifically, the connection relationship between the devices is: the outlet of the desorption desulfurization reactor 1 positioned at the bottom of the tower is connected with the inlet (positioned at the top of the tower) of the hydrogenation reactor 2; the outlet of the hydrogenation reactor 2 positioned at the bottom of the tower is connected with the inlet of the first gas-liquid separator 3; the outlet of the first gas-liquid separator 3 positioned at the bottom of the tower is connected with the inlet of the first rectifying tower 4; the outlet of the first rectifying tower 4 positioned on the measuring line is connected with the inlet of the second rectifying tower 5; the outlet of the second rectifying tower 5 positioned at the top of the tower is connected with the inlet of the third rectifying tower 6; the outlet of the third rectifying tower 6 positioned at the top of the tower is connected with the inlet of the orthosteric reactor 7; the outlet of the orthosteric reactor 7 positioned at the bottom of the tower is connected with the inlet of the second gas-liquid separator 8, the outlet of the second gas-liquid separator 8 positioned at the top of the tower and the outlet of the first gas-liquid separator 3 positioned at the top of the tower are respectively connected with the inlet of the circulating hydrogen compressor 9; the outlet of the recycle hydrogen compressor 9 is connected with the inlet of the hydrogenation reactor 2 and the inlet of the orthographic formation reactor 7 respectively.

The process flow of the process for producing solvent oil from the reformate raffinate used in examples 2a to 2e is shown in fig. 1, and specifically includes:

s1, conveying reformed raffinate oil to an adsorption reactor for desulfurization and dehydration;

S2, mixing the desulfurization and dehydration products with hydrogen, and feeding the mixture into a hydrogenation reactor for benzene removal and olefin removal, wherein a hydrogenation catalyst filled in the hydrogenation reactor is the hydrogenation catalyst provided by the invention;

S3, conveying the hydrogenation reaction product to a gas-liquid separator for gas-liquid separation, discharging the separated liquid product from the bottom of the gas-liquid separator, discharging the separated hydrogen-rich gas from the top of the gas-liquid separator, mixing the hydrogen-rich gas with new hydrogen, and boosting the pressure by a circulating hydrogen compressor to be used as circulating hydrogen;

s4, enabling the liquid product separated from the gas and the liquid to enter a first rectifying tower, wherein the top of the first rectifying tower is provided with a C5 component (pentane foaming agent), the lateral line of the first rectifying tower is provided with a C6 component, and the bottom of the first rectifying tower is provided with a C7 and above component (120 # solvent oil);

s5, C6 components coming out of the side line of the first rectifying tower enter a second rectifying tower for fractionation, the top of the tower is a mixture of normal hexane and isohexane, and the bottom of the tower is cyclohexane;

s6, the mixture of the n-hexane and the isohexane enters a third rectifying tower for fractionation, the isohexane is arranged at the top of the tower, and the n-hexane is arranged at the bottom of the tower;

S7, enabling the isohexane to enter a normal-structured reactor for reaction, enabling the obtained product to enter a separator for gas-liquid separation, enabling the product at the bottom of the separator to be high-concentration normal hexane, and then mixing the high-concentration normal hexane with the normal hexane product at the bottom of the third rectifying tower to further improve the normal hexane yield.

Table 1 shows the reformate feedstock compositions in examples 2a through 2e, comparative example 1, and comparative example 2 below.

TABLE 1

Example 2a

The embodiment provides a method for producing solvent oil by reforming raffinate oil, which comprises the following steps:

S1, taking reformed raffinate oil in the table 1 as a raw material, and allowing the reformed raffinate oil to enter an adsorption reactor for desulfurization and dehydration, wherein the catalyst is an industrial 5A molecular sieve adsorbent;

s2, mixing the desulfurization and dehydration products with hydrogen, and feeding the mixture into a hydrogenation reactor for benzene removal and olefin removal, wherein the catalyst is the high-nickel hydrogenation catalyst (Ni metal dispersity is more than 10%) prepared in the embodiment 1 a;

Adsorption desulfurization reaction conditions: normal temperature and pressure, airspeed of 3h ^-1, and product S is less than or equal to 1mg/kg;

S3, enabling a hydrogenation reactor product to enter a separator for gas-liquid separation, discharging a liquid product from the bottom of the separator, discharging hydrogen-rich gas from the top of the separator, mixing the hydrogen-rich gas with new hydrogen of a reforming device, and boosting the pressure of a circulating hydrogen compressor to be used as circulating hydrogen;

Reaction conditions of the fixed bed hydrogenation reactor: the reaction temperature is 120 ℃, the reaction pressure is 1.0MPa, and the hydrogen-oil volume ratio is 100: 1. the volume space velocity is 3.0h ^-1, and the benzene removal rate is 100 percent;

S4, enabling a liquid product at the bottom of the separator to enter a first rectifying tower, wherein the top of the first rectifying tower is a C5 component, a C6 component is discharged by a measuring line, and the bottom of the first rectifying tower is a C7 component or more;

First rectifying column operating conditions: the pressure in the tower is 0.1MPa, the temperature of the top of the tower is 35 ℃, the temperature of the bottom of the tower is 95 ℃, the C5 component of the top of the tower is pentane foaming agent, the C7 and above component of the bottom of the tower is 120# solvent oil, and the C6 component is discharged from the side line; bromine index determination is carried out by adopting a bromine valence bromine index determinator according to the microcoulomb titration principle, wherein the bromine index of a C5 component is 5.1mgBr/100g, and the bromine index of a C7 component and above component is 9.8mgBr/100g;

s5, enabling the C6 component coming out from the side line to enter a second rectifying tower for fractionation, wherein the top of the tower is a mixture of normal hexane and isohexane, and the bottom of the tower is cyclohexane;

second rectifying column operating conditions: the pressure at the top of the tower is 0.1MPa, the temperature of the top of the tower is 60 ℃, the temperature of the bottom of the tower is 72 ℃, and the purity of cyclohexane at the bottom of the tower is 97wt%;

third rectifying column operating conditions: the pressure at the top of the tower is 0.1MPa, the temperature of the top of the tower is 55 ℃, the bottom temperature is 65 ℃, the purity of the n-hexane at the bottom of the tower is 87%;

S7, enabling the top product isohexane to enter a normal structuring reactor for reaction, enabling the obtained product to enter a separator for gas-liquid separation, enabling the bottom product of the separator to be high-concentration normal hexane, and then mixing the high-concentration normal hexane with the normal hexane product at the bottom of the third rectifying tower to further improve the normal hexane yield.

The reaction conditions of the orthosteric reactor are as follows: the reaction temperature is 380 ℃, the reaction pressure is 2.0MPa, and the hydrogen-oil volume ratio is 200: 1. volume space velocity 1.5h ^-1, conversion 91%, selectivity >99wt%;

The catalyst used in the orthosteric reaction is a molecular sieve catalyst which consists of a molecular sieve carrier, a non-molecular sieve carrier and a third active metal component, wherein the molecular sieve carrier is a mercerized molecular sieve MOR, and the weight content of the molecular sieve carrier in the catalyst is 80%; the third active metal component is oxide of palladium, and the weight ratio of the third active metal component in the catalyst is 0.35%; the balance being alumina as a non-molecular sieve catalyst.

The product is obtained in this example: the purity of the n-hexane is 83wt percent, and the yield is 54 percent; the purity of cyclohexane is >97wt%; the C5 component meets the standard requirement of the pentane foaming agent; c7 and above components meet the standard requirements of 120# solvent oil.

The benzene removal rate, conversion rate, selectivity, purity and yield of each component in the product are all obtained by gas chromatography for analysis of the composition of raw materials and product families.

Example 2b

The raw materials, the process flow, the adsorption desulfurization reaction, the hydrogenation reaction and the orthographic structuring reaction were exactly the same as in example 2a, except that the operation conditions of the three rectifying towers were changed.

First rectifying column operating conditions: the pressure in the tower is 0.1MPa, the temperature of the tower top is 40 ℃, the temperature of the tower bottom is 102 ℃, the C5 component of the tower top is pentane foaming agent, the bromine index is 5.8mgBr/100g, the C7 and above component of the tower bottom is 120# solvent oil, and the bromine index is 8.3mgBr/100g;

Second rectifying column operating conditions: the pressure at the top of the tower is 0.1MPa, the temperature of the top of the tower is 66 ℃, the bottom temperature is 78 ℃, and the purity of cyclohexane at the bottom is 99wt%;

third rectifying column operating conditions: the pressure at the top of the tower is 0.1MPa, the temperature of the top of the tower is 60 ℃, the bottom temperature was 71 c, the purity of the n-hexane at the bottom of the tower is 91%;

the product is obtained in this example: n-hexane purity 88wt%, yield 53%; the purity of cyclohexane is >99wt%; the C5 component meets the standard requirement of the pentane foaming agent; c7 and above components meet the standard requirements of 120# solvent oil.

Example 2c

The raw materials, the process flow, the adsorption desulfurization reaction, the hydrogenation reaction and the orthographic structuring reaction conditions are exactly the same as in example 2b, except that the three rectifying towers are changed in operation conditions.

First rectifying column operating conditions: the pressure in the tower is 0.1MPa, the temperature of the tower top is 50 ℃, the temperature of the tower bottom is 110 ℃, the C5 component of the tower top is pentane foaming agent, the bromine index is 7.8mgBr/100g, the C7 and above component of the tower bottom is 120# solvent oil, and the bromine index is 8.9mgBr/100g;

second rectifying column operating conditions: the pressure at the top of the tower is 0.1MPa, the temperature of the top of the tower is 75 ℃, the bottom temperature is 90 ℃, and the purity of cyclohexane at the bottom is 98wt%;

third rectifying column operating conditions: the pressure at the top of the tower is 0.1MPa, the temperature of the top of the tower is 65 ℃, the bottom temperature was 78 c, the purity of the n-hexane at the bottom of the tower is 88%;

The product is obtained in this example: n-hexane purity 85wt%, yield 51%; purity of cyclohexane >98wt%; the C5 component meets the standard requirement of the pentane foaming agent; c7 and above components meet the standard requirements of 120# solvent oil.

Example 2d

wherein the raw materials, process flow, rectifying column operating conditions, adsorption desulfurization reaction and orthosteric reaction conditions are exactly the same as in example 2b, except for fixed bed hydrogenation reaction conditions.

Reaction conditions of the fixed bed hydrogenation reactor: the reaction temperature is 180 ℃, the reaction pressure is 1.0MPa, and the hydrogen-oil volume ratio is 100: 1. volume space velocity is 3.0h ^-1, benzene removal rate is 100%;

first rectifying column product: the bromine index of the C5 component at the top of the tower is 1.2mgBr/100g, the requirement of a pentane foaming agent is met, the bromine index of the C7 component at the bottom of the tower and above is 3.5mgBr/100g, and the requirement of 120# solvent oil is met;

The product is obtained: the purity of the n-hexane is 88wt percent, and the yield is 53 percent; the purity of cyclohexane is >99wt%; the C5 component meets the standard requirement of the pentane foaming agent; c7 and above components meet the standard requirements of 120# solvent oil.

Example 2e

Wherein the raw materials, the process flow, the operating conditions of the rectifying column, the adsorption desulfurization reaction and the fixed bed hydrogenation reaction are exactly the same as in example 2b, except for the orthosteric reaction conditions.

The reaction conditions of the orthosteric reactor are as follows: reaction temperature 300 ℃, reaction pressure 2.0MPa and hydrogen-oil volume ratio 200: 1. volume space velocity 1.5h ^-1, conversion 73%, selectivity >99wt%;

The product is obtained: the purity of the n-hexane is 82 weight percent, and the yield is 52 percent; the purity of cyclohexane is >99wt%; the C5 component meets the standard requirement of the pentane foaming agent; c7 and above components meet the standard requirements of 120# solvent oil.

Comparative example 1

The comparative example provides a method for producing solvent oil by reforming raffinate oil:

Reaction conditions of the fixed bed hydrogenation reactor: the reaction temperature is 80 ℃, the reaction pressure is 1.0MPa, and the hydrogen-oil volume ratio is 100: 1. volume space velocity is 3.0h ^-1, benzene removal rate is 83%;

First rectifying column product: the bromine index of the C5 component at the top of the tower is 121.2mgBr/100g, which does not meet the requirement of a pentane foaming agent, and the bromine index of the C7 component at the bottom of the tower is 133.7mgBr/100g, which does not meet the requirement of 120# solvent oil;

the comparative example gives the product: purity of n-hexane 76wt%, yield 53%; the purity of cyclohexane is >99w%.

Comparative example 2

This comparative example provides a process for the production of a solvent oil from the reforming of raffinate using the same starting materials as in example 2b, comprising:

1. The reforming raffinate oil raw material is subjected to benzene removal and olefin removal by a hydrogenation reactor, and the catalyst is a conventional hydrogenation olefin removal catalyst;

the reaction conditions of the hydrogenation reactor are as follows: the reaction temperature is 120 ℃, the reaction pressure is 1.0MPa, and the hydrogen-oil volume ratio is 100:

1. Volume space velocity is 3.0h ^-1, benzene removal rate is 82%;

The preparation method comprises the following steps: 82.9g of aluminum nitrate nonahydrate is weighed, and 240ml of aluminum nitrate nonahydrate is prepared into 1mol/L aluminum nitrate solution. 190ml of 1mol/L nickel nitrate solution, 30ml of 1mol/L magnesium nitrate solution and 170ml of 1mol/L aluminum nitrate solution are measured and uniformly mixed to obtain a metal mixed solution for later use. Weighing 1.5mol/L potassium carbonate as 1000ml of precipitant solution for later use. 40ml of purified water was added to the reaction vessel, and the reaction temperature was 50℃by heating at a stirring speed of 250 rmp. Under stirring, 35ml/min and 71ml/min of the metal mixed solution and the precipitant solution were added in parallel, respectively, and the pH was controlled to be constant at 9. After the completion of the dropwise addition of the solution, the mixture was aged at 50℃for 2 hours with stirring. The obtained material is filtered and washed until the filtrate is neutral. Drying the filter cake in an oven at 120 ℃ for 4 hours, grinding, roasting in a muffle furnace at 500 ℃ for 5 hours, and grinding to obtain the catalyst

2. The hydrogenation reaction product enters a separator to carry out gas-liquid separation, liquid products are discharged from the bottom, hydrogen-rich gas discharged from the top is mixed with new hydrogen of a reforming device, and the mixture is boosted by a circulating hydrogen compressor and then is used as circulating hydrogen;

3. The liquid product at the bottom of the separator enters a first rectifying tower for fractionation, C5 components are discharged from the top of the tower, C6 components are discharged from the side line, and C7 and above components are discharged from the bottom of the tower;

first rectifying column operating conditions: the pressure in the tower is 0.1MPa, the temperature of the tower top is 40 ℃, the temperature of the tower bottom is 102 ℃, the component C5 at the tower top is pentane foaming agent, the bromine index is 131.1mgBr/100g, the requirement of the pentane foaming agent is not met, the bromine index of the component C7 at the tower bottom and above is 138.9mgBr/100g, and the requirement of 120# solvent oil is not met;

4. The C6 component enters a second rectifying tower for fractionation, the top of the second rectifying tower is a mixture of normal hexane and isohexane, and the bottom of the second rectifying tower is cyclohexane;

Second rectifying column operating conditions: the pressure at the top of the tower is 0.1MPa, the temperature of the top of the tower is 66 ℃, the bottom temperature is 78 ℃, and the purity of cyclohexane at the bottom is 98wt%;

5. The mixture of normal hexane and isohexane enters a third rectifying tower for fractionation, the top of the third rectifying tower is isohexane, and the bottom of the third rectifying tower is normal hexane;

Third rectifying column operating conditions: the pressure at the top of the tower is 0.1MPa, the temperature of the top of the tower is 60 ℃, the bottom temperature was 71 ℃.

The comparative example gives the product: the purity of the n-hexane is 88wt percent, and the yield is 31 percent; the purity of the isohexane is 91wt% (which can be used as 6# solvent oil), and the purity of the cyclohexane is >99wt%.

Table 2 shows the results of evaluating the products of examples 2a to 2e and comparative examples 1 to 2.

TABLE 2

As can be seen from table 2, the evaluation results of examples 2a to 2e are significantly better than those of comparative examples 1 and 2. The reason is that:

In comparative example 2, a process flow of hydrogenation and then cutting is adopted, a conventional hydrogenation catalyst is used for treating the reformed raffinate, the hydrogenation performance of the catalyst is slightly poor, the removal rate of olefin and benzene is low, the standard requirement of bromine index of a pentane foaming agent cannot be met by the presence of olefin in a C5 component, and the sulfur content requirement of 120# solvent oil cannot be met by the components C7 and above because sulfolane is not removed; meanwhile, as the isohexane orthosteric reaction unit is not arranged, the isohexane cannot be converted into the normal hexane, the yield of the normal hexane is reduced, and the high-efficiency utilization of the reformed raffinate oil cannot be realized.

In contrast, the process flows of examples 2a to 2e of the present invention are used for treating the reformed raffinate, and by adding the adsorption desulfurization unit and the orthosteric reaction unit, the reformed raffinate is efficiently utilized.

Wherein, in the embodiment 2a to the embodiment 2c, the regulation and control of the purity and the yield of the n-hexane can be realized by adjusting the operation temperature of the top and the bottom of the rectifying tower; example 2b and example 2d can achieve different degrees of benzene and olefin removal by changing the operating temperature of the hydrogenation catalyst, in contrast to the comparative example 1 where the operating temperature of the hydrogenation catalyst did not reach 90-200 ℃, the benzene and olefin removal was incomplete; examples 2b and 2e can realize different degrees of normal formation of isohexane to form n-hexane by changing normal formation reaction conditions, and the purity of the n-hexane is adjusted.

The results show that the method for producing the solvent oil from the reformed raffinate oil is suitable for separating and utilizing the reformed raffinate oil with higher benzene and olefin content, and the produced solvent oil has high purity and yield and good utilization effect, and can effectively improve the added value of the reformed raffinate oil.

Claims

1. A method for preparing a hydrogenation catalyst, wherein the method comprises:

the temperature of the slurry of the carrier precursor reaches the precipitation reaction temperature, and the pH value of the slurry of the carrier precursor is regulated to reach the alkaline pH value, so that a first mixed system is obtained;

Adding a first active metal component precursor, a second active metal component precursor and a precipitant into the first mixed system in parallel flow, and keeping the pH value of the system at an alkaline pH value to obtain a second mixed system;

Aging the second mixed system, and roasting the aging product to obtain the hydrogenation catalyst;

The hydrogenation catalyst comprises a carrier, and a first active metal component and a second active metal component which are loaded on the carrier, wherein the metal element of the first active metal component comprises nickel, the weight content of the first active metal component in the catalyst is 40-70% based on the total weight of the catalyst, the weight content of the carrier in the catalyst is 20-59%, and the balance is the second active metal component.

2. The method according to claim 1, wherein the alkaline pH is 8-11.

3. The preparation method according to claim 1, wherein the precipitation reaction temperature is 40-90 ℃.

4. The production method according to claim 1, wherein the temperature of the aging treatment is 40 to 90 ℃.

5. The production method according to claim 1, wherein the baking temperature is 300 to 600 ℃.

6. The production method of a hydrogenation catalyst according to claim 1, wherein the first active metal component precursor is a compound of a metal element in the first active metal component;

The second active metal component precursor is a compound of a metal element in the second active metal component;

The precipitant includes an alkaline precipitant.

7. The method for producing a hydrogenation catalyst according to claim 6, wherein the first active metal component precursor comprises a metal salt of a metal element in the first active metal component.

8. The method for producing a hydrogenation catalyst according to claim 6, wherein the second active metal component precursor comprises a metal salt of a metal element in the second active metal component.

9. The method for preparing a hydrogenation catalyst according to claim 6, wherein the precipitant comprises a soluble carbonate.

10. A hydrogenation catalyst obtainable by the process for the preparation of a hydrogenation catalyst according to any one of claims 1 to 9.

11. The hydrogenation catalyst according to claim 10, wherein the dispersity of Ni metal in the hydrogenation catalyst is 10% or more;

The metal element of the second active metal component comprises one or more than two of copper, lanthanum and magnesium;

The carrier comprises one of alumina, silicon oxide, titanium oxide and cerium oxide, and/or a mixture of at least one of the silicon oxide, the titanium oxide and the cerium oxide and the alumina.

12. The hydrogenation catalyst of claim 10, wherein the hydrogenation catalyst has a specific surface area of 200-500m ²/g and a pore volume of 0.3-0.6cm ³/g.

13. A process for reforming raffinate oil to produce solvent oil, the process comprising:

S2, catalyzing the reformed raffinate oil subjected to desulfurization and dehydration by using a hydrogenation catalyst to carry out hydrogenation and debenzolization and olefin removal reactions to obtain a reaction product;

s6, rectifying the mixture of the normal hexane and the isohexane, and respectively distilling off the isohexane and the normal hexane to finish the production of solvent oil;

Wherein the hydrogenation catalyst comprises the hydrogenation catalyst of any one of claims 10-12.

14. The method of claim 13, wherein in S2, the hydrodebenzene, dealkenation reaction conditions are: the temperature is 90-200 ℃, the reaction pressure is 1.0-3.0MPa, and the hydrogen-oil volume ratio is 100-300: 1. volume airspeed 1.0-3.0h ^-1;

s4, rectifying under the pressure of 0-0.2MPa;

S5, rectifying under the pressure of 0-0.1MPa;

in S6, the pressure of the rectification is 0-0.1MPa.

15. The process according to claim 14, wherein in S2, the benzene content in the reaction product is 0.01% or less and the bromine index is 50mgBr/100g or less.

16. The process according to claim 14, wherein the rectification of S4 is carried out in a rectification column having a top temperature of 30-50 ℃ and a bottom temperature of 90-110 ℃.

17. The process according to claim 14, wherein the rectification of S5 is carried out in a rectification column having a top temperature of 55-75 ℃ and a bottom temperature of 70-90 ℃.

18. The process according to claim 14, wherein the rectification of S6 is carried out in a rectification column having a top temperature of 50-70 ℃ and a bottom temperature of 60-80 ℃.

19. The process of claim 13 wherein the non-desulfurized, dehydrated reformate has a sulfur content of less than 20mg/kg, a benzene content of 3-5% by volume, an olefin content of 1-3% by volume, and a total content of n-hexane and isohexane of greater than 50% by volume.

20. The method of claim 19, wherein the method further comprises: s1, desulfurizing and dehydrating the heavy raffinate oil, wherein the S content of the product is less than or equal to 1mg/kg.

21. The method of claim 20, wherein in S1, the space velocity during the desulfurizing and dehydrating process is 1-10h ^-1.

22. The method of claim 20, wherein in S1, the adsorbent used for desulfurization and dehydration comprises a molecular sieve adsorbent.

23. The method of claim 22, wherein the molecular sieve adsorbent comprises one or a combination of more than two of a 3A molecular sieve, a 4A molecular sieve, a 5A molecular sieve, and a 13A molecular sieve.

24. The method of claim 22, wherein the molecular sieve adsorbent has a specific surface area of 400-700m ²/g and a pore volume of 0.05-0.30cm ³/g.

25. The method according to any one of claims 13-24, wherein the method further comprises: s7, converting the isohexane obtained in the step S6 into normal hexane through a normal structuring reaction.

26. The method of claim 25, wherein the conditions of the orthosteric reaction are: the reaction temperature is 200-400 ℃, the reaction pressure is 1.0-3.0MPa, and the hydrogen-oil volume ratio is 100-300: 1. the volume space velocity is 1.0-3.0h ^-1.

27. The method of claim 25, wherein the isohexane has a purity of 90% or greater.

28. The method of claim 27, wherein the isohexane has a purity of 95% or greater.

29. The method of claim 25, wherein the catalyst of the orthographic reaction comprises a molecular sieve catalyst comprising a third active metal component, a molecular sieve support, and a non-molecular sieve support.

30. The method of claim 29, wherein the molecular sieve support comprises one or a combination of two or more of MOR, MCM-41, ZSM-22, SAPO-11.

31. The method of claim 29, wherein the metallic element in the third active metal component comprises an element from group VIB and/or VIII.

32. The method of claim 29, wherein the non-molecular sieve support comprises alumina.

33. The process of claim 29 wherein the weight of the molecular sieve support is from 10 to 80% of the total weight of the catalyst, the weight of the third active metal component is from 0.01 to 5% of the total weight of the catalyst, and the balance is a non-metallic support, based on 100% of the total weight of the molecular sieve catalyst.

34. The method of claim 33, wherein the weight of the third active metal component is from 0.1 to 0.5% of the total weight of the catalyst, based on 100% of the total weight of the molecular sieve catalyst.

35. The method of claim 33, wherein the weight of the third active metal component is 0.1-0.5% of the total weight of the catalyst.

36. A system for reforming raffinate oil to produce solvent oil, the system being capable of carrying out the method for reforming raffinate oil to produce solvent oil of any one of claims 13 to 35.

37. The system for producing solvent oil from reformed raffinate oil of claim 36 wherein the system for producing solvent oil from reformed raffinate oil comprises a desorption desulfurization reactor, a hydrogenation reactor, a gas-liquid separator, a first rectifying tower, a second rectifying tower and a third rectifying tower which are connected in sequence.

38. The system for producing solvent oil from reformed raffinate oil of claim 37 wherein the system for producing solvent oil from reformed raffinate oil further comprises a orthosteric reactor, an inlet of the orthosteric reactor being connected to an outlet of the third rectifying column, and an outlet of the orthosteric reactor being connected to an inlet of the gas-liquid separator.