CN116023211A - Method and system for preparing hydrogenated terphenyl and obtained hydrogenated terphenyl - Google Patents

Abstract

The invention discloses a method and a system for preparing hydrogenated terphenyl and the obtained hydrogenated terphenyl, wherein the method comprises the following steps: with hydrogen and from dec ₆ Column C ₆ The components are raw materials for N-stage reaction, N is more than or equal to 2, and reaction products enter into C removal after the N-stage reaction is finished ₆ A tower; wherein, at the dec ₆ In the tower, fresh benzene is taken as a feed, and the C is laterally extracted ₆ The components are extracted from the tower top materials, and the product containing hydrogenated terphenyl is extracted from the tower bottom. The invention takes cheap benzene and hydrogen as raw materials, and can realize the preparation of hydrogenated terphenyl at the temperature of 130-280 ℃ (preferably 150-220 ℃) and the pressure of 0.5-2.5 MPaG (preferably 1.0-2.0 MPaG), thereby avoiding the high temperature of benzeneAnd (3) a cracking operation.

Description

Method and system for preparing hydrogenated terphenyl and obtained hydrogenated terphenyl

Technical Field

The invention relates to preparation of hydrogenated terphenyl, in particular to a method and a system for preparing hydrogenated terphenyl and the obtained hydrogenated terphenyl.

Background

The heat conducting oil is used as an effective heat transfer medium and is widely applied to petroleum and petrochemical industry, manufacturing industry, food industry and the like. At present, domestic heat conduction oil mainly comprises mineral oil type heat conduction oil with the applicable temperature below 300 ℃. With the increasingly strict environmental protection requirements in China and the continuous development of new industries such as photovoltaic power generation, energy storage, polyester and the like and new technologies, the requirements of synthetic high-temperature heat conduction oil such as L-QD340 with long service life and higher use temperature (more than or equal to 330 ℃) are greatly increased. Hydrogenated terphenyl is the main base oil raw material of L-QD340, and the product is mainly imported at present, so that the domestic productivity is extremely limited, and the urgent market demand at present cannot be met.

Patent CN109053355a relates to a method for purifying biphenyl by continuous rectification, wherein raw material liquid is fed into a benzene removal tower, a biphenyl tower and a terphenyl tower which are connected in series, a small amount of hydrogen, benzene and biphenyl are extracted from the top of the benzene removal tower, and the rest of biphenyl, terphenyl mixture and trinitro tar are extracted from the tower bottom and enter the biphenyl tower; a small amount of biphenyl mixture containing benzene is extracted from the top of the biphenyl tower and returned to the reaction liquid storage tank, high-purity biphenyl is extracted from the side line, and the terphenyl and trinity tar mixture extracted from the tower bottom enters the terphenyl tower; a small amount of biphenyl mixture containing terphenyl is extracted from the top of the terphenyl tower and returned to the feed inlet of the biphenyl tower, high-purity terphenyl mixture is extracted from the side line, and heavy component trinitro tar is extracted from the tower bottom. By arranging the side line extraction, the high-purity biphenyl product with the mass fraction of more than 99.9 percent can be obtained while the method is suitable for large-scale production and has high operation elasticity, and compared with the traditional batch process, the energy consumption of the unit product can be reduced by 30-40 percent, and the method has extremely high economic value.

The patent CN103804114A is characterized in that benzene is cracked at high temperature to obtain a reaction product containing benzene, biphenyl and terphenyl, the reaction product is cooled and rectified to obtain terphenyl, and the terphenyl is hydrogenated in a hydrogenation autoclave to obtain hydrogenated terphenyl. The above processes all give hydrogenated terphenyl products, but the starting material is still derived from the cleavage of benzene at high temperatures.

However, in the prior art, the traditional industrial production method of hydrogenated terphenyl is a benzene high-temperature cracking hydrogenation method, which is a byproduct of producing biphenyl, and the method has high cracking temperature (more than or equal to 800 ℃), large energy consumption and low total yield (4%).

Disclosure of Invention

In order to overcome the problems in the prior art, the invention provides a method and a system for preparing hydrogenated terphenyl and the obtained hydrogenated terphenyl, which mainly solve the problems of high raw material cost, high energy consumption and low total yield in the prior art. The invention adopts N-level reactors to react in series, and the feed of the 1 st-level reactor is hydrogen and the feed comes from C removal ₆ Column C ₆ The components, the feed of each subsequent stage reactor is hydrogen and the reaction product from the previous stage reactor; the reactor is provided with an inter-stage cooler, and the cooling medium is from the C removal ₆ Column C ₆ The components, the reaction product of the N-stage reactor and fresh benzene enter the C-removing reactor from different positions respectively ₆ Column, C separated by column ₆ The components are returned to the 1 st stage reactor step by step through a cooler, and the bottom stream enters the technical scheme of recovering hydrogenated terphenyl in the subsequent working section, so that the problem is well solved, and the method can be used in the industrial production of hydrogenated terphenyl.

It is an object of the present invention to provide a method for preparing hydrogenated terphenyl, comprising: with hydrogen and from dec ₆ Column C ₆ The components are raw materials for N-stage reaction, N is more than or equal to 2, and the reaction product after the N-stage reaction (i.e. the last stage) is finished enters the C removal device ₆ A tower; wherein, at the dec ₆ In the tower, fresh benzene is taken as a feed, and the C is laterally extracted ₆ The components are extracted from the tower top materials, and the product containing hydrogenated terphenyl is extracted from the tower bottom.

Wherein along hydrogen and C ₆ Of componentsIn a flow mode, the 1 st stage reaction, …, the (N-1) th stage reaction and the N-th stage reaction are sequentially carried out. At the step of removing C ₆ Heavy components (including C) in the N-stage reaction product in the column ₆ Above contains no C ₆ Component (C) of (C) ₁₂ And C ₁₈ ) Extracted from the tower kettle, and C ₆ The above components (including unreacted benzene and cyclohexane) are subjected to C removal ₆ An upper portion of the column. Thus, the C ₆ The component contains benzene and cyclohexane, wherein a part of benzene is fresh benzene, and a part of benzene is unreacted benzene and byproduct cyclohexane in the N-stage reaction product (the cyclohexane is recycled to the subsequent reaction and has no influence on the reaction).

In the invention, hydrogen and benzene are directly reacted to form hydrogenated terphenyl, so that a high-temperature cracking mode is avoided.

In a preferred embodiment, the process comprises N-stage reactions, the hydrogen being fed separately to each of the stages in N strands from the decarbonization ₆ Column C ₆ The components enter each section in turn for reaction in one strand; that is, the feed for the stage 1 reaction is hydrogen and the C ₆ The components, the feed to each subsequent stage reaction is hydrogen and the reaction product from the previous stage reaction.

In a further preferred embodiment, n=2 to 6, preferably 2 to 4, for example n=2, 3, 4, 5 or 6.

In a preferred embodiment, the reaction product of the previous stage reaction is subjected to a cooling treatment before it enters the next stage reaction.

The reaction of hydrogen and benzene in the invention is exothermic, the temperature of the reaction product is higher after each section of reaction is finished, and if the reaction product directly enters the next section of reaction without cooling, the temperature of the subsequent reaction is always increased, thereby influencing the performance of the catalyst.

In a further preferred embodiment, to come from dec ₆ Said C of the column ₆ The components are cooling medium.

In a still further preferred embodiment, in said C ₆ The components are reacted with the (N-1) th section and the (N-2) th section in turn before entering each section reactionThe reaction products of the reaction of the first section and … are subjected to heat exchange treatment, and then enter each section of reaction in turn after heat exchange.

Wherein C is ₆ The components need to be heated to a temperature close to the reaction temperature before entering the stage 1 reaction, if the means disclosed in the prior art are adopted, the electric heater is adopted to perform the reaction on the component C ₆ The components are heated and then enter the 1 st stage reaction. However, in this application, C is used ₆ The reaction products between the components and the segments exchange heat, on the one hand C ₆ The components can be heated to the required temperature, and on the other hand, the reaction products between the sections are cooled, so that the energy is greatly saved.

In a preferred embodiment, the N-stage reactions are each independently carried out in the presence of a catalyst.

In a further preferred embodiment, the catalyst is a metal supported bi-functional solid acid hydroalkylation catalyst.

In the present invention, the metal supported bifunctional solid acid hydroalkylation catalyst may be selected from those disclosed in the prior art, preferably but not limited to: the catalyst comprises a carrier, a binder and an active component, wherein the carrier is an organosilicon microporous zeolite, the active component comprises a combination of a noble metal active component (preferably at least one selected from palladium, ruthenium or platinum) and a non-noble metal (preferably selected from nickel, copper and cobalt), and the binder is at least one selected from zinc oxide, aluminum oxide or titanium oxide.

Preferably, the carrier is present in an amount of 30 to 90wt%, the binder is present in an amount of 8 to 60wt%, the noble metal active component is present in an amount of 0.05 to 5wt%, and the non-noble metal active component is present in an amount of 10 to 60wt%, based on 100wt% of the total weight of the catalyst.

For example, the carrier may be present in an amount of 30wt%, 40wt%, 50wt%, 60wt%, 70wt%, 80wt%, or 90wt%, the binder may be present in an amount of 8wt%, 10wt%, 20wt%, 30wt%, 40wt%, 50wt%, or 60wt%, the noble metal active component may be present in an amount of 0.05wt%, 0.1wt%, 0.5wt%, 1wt%, 2wt%, 3wt%, 4wt%, or 5wt%, and the non-noble metal active component may be present in an amount of 10wt%, 25wt%, 30wt%, 35wt%, 40wt%, 45wt%, 50wt%, 55wt%, or 60wt%, based on 100wt% of the total weight of the catalyst.

In the present invention, the silicone microporous zeolite can be purchased directly or prepared by methods disclosed in the prior art. Preferably, the silicone microporous zeolite has a composition in the following molar relationship: (1/n) A12O3 is SiO2 (m/n) R, wherein n=5 to 250, m=0.01 to 50, and R is at least one of alkyl or phenyl with 1 to 8 carbon atoms; si of the organosilicon microporous zeolite ²⁹ The NMR solid nuclear magnetic spectrum at least comprises one Si between-80 and +50ppm ²⁹ Nuclear magnetic resonance spectrum peaks; the X-ray diffraction pattern of the silicone microporous zeolite has d-spacing maxima at 12.4±0.2, 11.0±0.3,9.3±0.3,6.8±0.2,6.1±0.2,5.5±0.2,4.4±0.2,4.0±0.2, and 3.4±0.1 angstrom.

In the present invention, preferably, the catalyst is prepared as follows: firstly grinding the organic silicon microporous zeolite into powder to prepare mixed solution containing noble metal active components and non-noble metal active components. Spraying the prepared mixed solution on the organic silicon microporous zeolite powder, and continuously stirring the organic silicon microporous zeolite in the spraying process. Air-drying at normal temperature and pressure for 1-20 hours, drying at 100-150 ℃ for 1-20 hours, and grinding into powder. Mixing the powder with adhesive, kneading to form, drying, roasting at 400-550 deg.c in air for 1-10 hr, exchanging in ammonium salt solution, washing, stoving and roasting at 480 deg.c in air for 5 hr to obtain the catalyst product.

In a preferred embodiment, the N-stage reactions are each independently carried out in a reactor, preferably the reactor is an adiabatic fixed bed reactor, more preferably each independently packed with the catalyst in each reactor.

In a preferred embodiment, the total amount of hydrogen is combined with C ₆ The molar ratio of benzene in the component (A) is 1 (1.1-6), preferably 1 (1.5-4.5).

For example, hydrogen and C ₆ The molar ratio of benzene in the components is 1:1.1, 1:1.2, 1:1.5, 1:21:2.5, 1:3, 1:3.5, 1:4, 1:4.5, 1:5, 1:5.5, or 1:6.

Wherein if the total amount of hydrogen is equal to C ₆ The molar ratio of benzene in the component is too high, so that the energy consumption is high, and too low, so that the service life of the catalyst is greatly reduced.

In a further preferred embodiment, the hydrogen feed is the same or different, preferably the same, in each reaction stage.

In a preferred embodiment, the hydrogen is reacted with the C before entering each stage of reaction ₆ The components or the reaction products of the hydrogen and the previous reaction are mixed in a mixer.

Wherein hydrogen is reacted with (after heat exchange) C prior to stage 1 reaction ₆ The components are firstly mixed in a mixer; before the 2 nd to N th stage reactions, the hydrogen and the reaction product of the previous stage reaction (after heat exchange) are mixed in a mixer.

In a further preferred embodiment, the distance between each stage of mixer and its corresponding feed location for each stage of reaction is greater than or equal to 0.5 meter, preferably greater than or equal to 1 meter.

Wherein, the materials can be controlled to be uniformly mixed, especially the hydrogen and the benzene are uniformly mixed.

In a preferred embodiment, the reaction inlet temperature of each stage reaction is independently 130 to 280℃and the pressure of each stage reaction is independently 0.5 to 2.5MPaG.

In a further preferred embodiment, the reaction inlet temperature of each stage reaction is independently from 150 to 220℃and the pressure of each stage reaction is independently from 1.0 to 2.0MPaG.

In a still further preferred embodiment, the reaction inlet temperature of each stage reaction is independently from 150 to 180℃and the pressure of each stage reaction is independently from 1.2 to 1.7MPaG.

For example, the reaction inlet temperature of each stage reaction is 150 ℃, 160 ℃, 170 ℃, 180 ℃, 200 ℃, or 220 ℃, and the pressure of each stage reaction is 1.0MPaG, 1.2MPaG, 1.4MPaG, 1.6MPaG, 1.8MPaG, or 2.0MPaG, respectively.

In a preferred embodimentWherein, in the dec ₆ The reaction product feed location for the N-stage reaction is below the feed location for fresh benzene in the column.

In a further preferred embodiment, the reaction product of the N-stage reaction is removed from the C ₆ The lower part of the tower is introduced, and the fresh benzene is removed from C ₆ The upper part of the column enters.

Wherein, the reaction product of the N-stage reaction can contain unreacted benzene, target product hydrogenated terphenyl, byproduct cyclohexane and byproduct C ₁₂ Components and C other than hydrogenated terphenyl ₁₈ The components are as follows. The inventors have found through a number of experiments that when the reaction product of the N-stage reaction is reacted in the said C removal ₆ When the lower part of the tower is fed and the C6 removing tower is controlled within a proper temperature and pressure range, unreacted benzene and byproduct cyclohexane can enter the C removing tower ₆ The upper part of the tower is laterally extracted together with fresh benzene, and the target product hydrogenated terphenyl and byproduct C ₁₂ Components and C other than hydrogenated terphenyl ₁₈ The components are extracted from the tower kettle. Thus, not only unreacted benzene and byproduct cyclohexane in the product are separated, but also the unreacted benzene can be further recycled.

In a still further preferred embodiment, in said dec ₆ In the column, the fresh benzene is fed at a position C ₆ Below the side offtake position of the components.

Wherein fresh benzene rises from the feed position to the side offtake and is side offtake. 1000-2000 ppm water (which has great influence on my catalyst) and other light impurities in fresh benzene, through C removal ₆ Column treatment, removal of C from fresh benzene water and light component impurities ₆ The top of the column being taken off, i.e. the said C removal ₆ The tower achieves the purification of fresh benzene. Thus, the side-draw C6 component is free of water and other light components.

In a preferred embodiment, the dec ₆ The operating temperature of the tower is 150-250 ℃ and the operating pressure is 20-200 kPaA.

In a further preferred embodiment, the dec ₆ The operating temperature of the tower is 180-220 ℃ and the operating pressure is 80-140 kPaA.

For example, the dec ₆ The operating temperature of the column is 150 ℃, 180 ℃, 200 ℃, 220 ℃, 240 ℃ or 250 ℃ and the operating pressure is 20kPaA, 50kPaA, 80kPaA, 100kPaA, 120kPaA, 150kPaA, 180kPaA or 200kPaA.

In the present invention, the heavy ends removal column is preferably a tray column.

In a preferred embodiment, the catalyst is removed from the catalyst ₆ C from the side of the column ₆ The component temperature is 60 to 120 ℃, for example 60 ℃, 70 ℃, 80 ℃, 90 ℃, 100 ℃, 110 ℃ or 120 ℃.

In the present invention, the C is removed from the catalyst ₆ The materials extracted from the tower bottom enter a subsequent working section to recycle the hydrogenated terphenyl (which is only disclosed in the prior art), and preferably, the subsequent working section comprises the hydrogenated terphenyl vacuum rectification and the heat recovery.

In a preferred embodiment, the method is performed as follows: adopting N sections of serial reactions, wherein N is an integer greater than or equal to 2; the reaction process comprises the following steps:

1) Hydrogen and C ₆ The components are mixed and then contact with a catalyst to enter a1 st stage reaction to generate a1 st stage material flow;

2) Grade I stream C ₆ The components are cooled, mixed with hydrogen and contacted with a catalyst to enter a 2 nd stage reaction to generate a 2 nd stage material flow;

3) When N is equal to 2, the stage II stream goes to stage 4), when N is greater than 2, the stage 2 stream C ₆ Cooling the components, mixing with hydrogen, contacting with a catalyst, entering into a 3 rd-stage reaction to generate a 3 rd-stage material flow, repeating the step 3) until the reacted material enters into an N-stage reaction to generate an N-stage material flow;

4) The Nth grade material flow and fresh benzene enter a C removing position ₆ A tower;

5)C ₆ component C removal ₆ The side line of the tower is extracted, the reaction products of each section of reaction are subjected to gradual heat exchange and then returned to the 1 st section of reaction, and the C is removed ₆ And (3) stripping tower top materials, and feeding tower bottom materials into a subsequent working section to recover hydrogenated terphenyl.

Second object of the present inventionIt is an object of the present invention to provide a system for producing hydrogenated terphenyl, preferably for carrying out one of the processes of the object of the present invention, comprising N-stage series reactors and a dec ₆ The tower, wherein, along the material flow direction, is 1 st level reactor, …, (N-1) th level reactor and N level reactor in proper order, and N is greater than or equal to 2.

In a preferred embodiment, the N-stage series reactor is N reactors in series, n=2 to 6, preferably 2 to 4, for example n=2, 3, 4, 5 or 6.

In a preferred embodiment, a feed port is provided in the upper part or top of each stage reactor and a discharge port is provided in the lower part or bottom of each stage reactor.

In a further preferred embodiment, the outlet of the upper stage reactor is connected to the inlet of the lower stage reactor.

In a preferred embodiment, in said dec ₆ The tower is provided with a reaction material inlet, a fresh benzene inlet, a side line outlet, an outer tower top outlet and an outer tower kettle outlet.

In a further preferred embodiment, in said dec ₆ On the tower, the reaction material inlet is arranged below the fresh benzene inlet, and the side-line extraction outlet is arranged above the fresh benzene inlet.

In a still further preferred embodiment, in said dec ₆ On the tower, the reaction material inlet is arranged on the C-removing device ₆ The lower part of the tower, the fresh benzene inlet and the side-line extraction outlet are arranged at the C-removing part ₆ The upper part of the tower and the side-draw outlet are arranged above the fresh benzene inlet.

In a preferred embodiment, the removal of C ₆ The reaction material inlet of the tower is connected with the discharge port of the Nth stage reactor; and/or take off C ₆ The tower top outer extraction port of the tower is arranged at the tower top; and/or take off C ₆ The tower kettle outer discharge and extraction outlet of the tower is arranged at the tower kettle.

In a preferred embodiment, at least one cooler, i.e. at least (N-1) coolers in total, are each provided independently between adjacent two-stage reactors.

In a further preferred embodiment, the removal of C ₆ The side offtake of the tower, all coolers and the feed inlet of the 1 st stage reactor are connected through pipelines in sequence; preferably, the removal of C is performed in a direction opposite to the flow of material between the reactors of the stages ₆ The side offtake of the tower is connected with each cooler in turn and then connected with the feed inlet of the 1 st stage reactor.

In this way, the material withdrawn from the side offtake passes sequentially through the coolers in a direction opposite to the flow of material between the reactors of the stages and finally into the reactor of stage 1.

In a preferred embodiment, mixers are provided separately before the feed inlet of each stage of the reactor for mixing hydrogen with C ₆ The components or the reaction products of the previous stage reactor are uniformly mixed.

In a further preferred embodiment, the distance between each mixer and the feed inlet of its corresponding reactor of each stage is greater than or equal to 0.5 meter, preferably greater than or equal to 1 meter.

It is a further object of the present invention to provide hydrogenated terphenyl obtainable by the process or the system of the present invention.

The endpoints of the ranges and any values disclosed in the present invention are not limited to the precise range or value, and the range or value should be understood to include values close to the range or value. For numerical ranges, one or more new numerical ranges may be found between the endpoints of each range, between the endpoint of each range and the individual point value, and between the individual point value, in combination with each other, and are to be considered as specifically disclosed herein. In the following, the individual technical solutions can in principle be combined with one another to give new technical solutions, which should also be regarded as specifically disclosed herein.

Compared with the prior art, the invention has the following beneficial effects:

(1) The preparation of hydrogenated terphenyl can be realized by taking cheap benzene and hydrogen as raw materials at the temperature of 130-280 ℃ (preferably 150-220 ℃) and the pressure of 0.5-2.5 MPaG (preferably 1.0-2.0 MPaG), so that the benzene pyrolysis operation is avoided;

(2) In C ₆ The components are used as cooling medium to realize accurate control of the reaction temperature and effectively utilize the reaction heat;

(3) The total conversion rate of benzene is higher than 78%, and the total yield of hydrogenated terphenyl is higher than 70%;

(4) Compared with the traditional method, the method has the advantages of low-cost and easily-obtained raw materials, mild reaction conditions, easy control of heat removal, low energy consumption, high selectivity of target products and the like.

Drawings

Fig. 1 shows a schematic diagram of an embodiment of the system according to the invention.

Description of the drawings:

r-101-1 st stage reactor, R-102-2 nd stage reactor, R-103-3 rd stage reactor, M-dec ₆ Tower, E-101-stage 1 cooler, E-102-stage 2 cooler, 1-hydrogen, 2-stage 1 reaction product, 3-stage 2 reaction product, 4-stage 3 reaction product, 5-fresh benzene, 6-C ₆ The components, 7-tower top materials and 8-tower bottom materials.

Detailed Description

The present invention is described in detail below with reference to specific embodiments, and it should be noted that the following embodiments are only for further description of the present invention and should not be construed as limiting the scope of the present invention, and some insubstantial modifications and adjustments of the present invention by those skilled in the art from the present disclosure are still within the scope of the present invention.

In addition, the specific features described in the following embodiments may be combined in any suitable manner without contradiction. The various possible combinations of the invention are not described in detail in order to avoid unnecessary repetition.

In addition, any combination of the various embodiments of the present invention can be made, so long as the concept of the present invention is not deviated, and the technical solution formed thereby is a part of the original disclosure of the present specification, and also falls within the protection scope of the present invention.

The catalysts used in the examples and comparative examples were prepared as follows:

(1) Synthesis of organosilicon microporous zeolite: 3.0g of alumina is dissolved in 450g of water, 16.0g of sodium hydroxide is added to dissolve the alumina, 34.7g of hexamethyleneimine is added under the condition of stirring, 60g of solid silica and 5.9g of dimethyl diethoxysilane are added, and the material ratio (molar ratio) of reactants is as follows: siO (SiO) ₂ /A1 ₂ O ₃ ＝30、NaOH/SiO ₂ =0.2, dimethyldiethoxysilane/SiO ₂ =0.04, hexamethyleneimine/SiO ₂ ＝0.35、H ₂ O/SiO ₂ =25; after the reaction mixture was stirred uniformly, it was transferred to a stainless steel reaction vessel and crystallized at 145℃for 70 hours with stirring. Taking out, filtering, washing and drying. Chemical analysis to obtain SiO ₂ /A1 ₂ O ₃ The molar ratio was 30.1.

(2) Synthesizing the catalyst: 40ml of a mixed solution of ruthenium chloride and copper chloride was prepared, wherein the Ru content in the solution was 0.25g, cuCl ₂ The content of (C) was 9.95g, and the molar ratio of Ru to Cu was 1:30. 32.50g of organic silicon microporous zeolite powder is taken, and the prepared mixed solution is sprayed on the organic silicon microporous zeolite powder, and the organic silicon microporous zeolite is continuously stirred in the spraying process. Air-drying at normal temperature and pressure for 10 hours, oven-drying at 120deg.C for 10 hours, and grinding into powder. 8.98g of alumina was added and mixed, and a dilute nitric acid solution was added thereto for kneading and extrusion to form a bar having a diameter of 1.6X2 mm. After drying, the mixture was calcined at 550℃for 5 hours, then exchanged with 1M ammonium nitrate 5 times, filtered and dried. Drying at 120 deg.c for 12 hr and roasting at 480 deg.c for 6 hr to prepare the required catalyst. Wherein, in weight percent, ru is 0.52%, cu is 12.5%, binder is 18.7%, organosilicon microporous zeolite is 67.7%, and other impurities.

[ example 1 ]

Certain 1 ten thousand tons/year benzene hydroalkylation hydrogen production terphenyl device adopts the process technology of figure 1, but the reactor is four-stage, the four-stage reactor is an adiabatic fixed bed reactor, the fresh benzene flow is 1800kg/h, the hydrogen flow of each section is 14kg/h, and the self-stripping C is realized ₆ C extracted from tower ₆ The flow rate of the components is 8000kg/h.

C ₆ The components pass through an interstage cooler step by step, are mixed with 1 st hydrogen in a1 st stage mixer, enter a1 st stage reactor under the conditions of the temperature of 150 ℃ and the pressure of 1.7MPaG, the temperature of the outlet gas of the 1 st stage reactor is 184 ℃, and the outlet gas passes through C ₆ Cooling the components, mixing with 2 nd hydrogen in a 2 nd stage mixer, introducing into a 2 nd stage reactor under the conditions of 150 ℃ and 1.65MPaG, wherein the temperature of the outlet gas of the 2 nd stage reactor is 187 ℃, and the outlet gas passes through C ₆ Cooling the components, mixing with 3 rd hydrogen in a 3 rd stage mixer, introducing into a 3 rd stage reactor under the conditions of the temperature of 151 ℃ and the pressure of 1.6MPaG, wherein the temperature of the outlet gas of the 3 rd stage reactor is 188 ℃, and the outlet gas passes through C ₆ Cooling the components, mixing with 4 th hydrogen in 4 th stage mixer, introducing into 4 th stage reactor at 152 deg.C under 1.55MPaG pressure, introducing into C-removing reactor at 188 deg.C under 1.5MPaG ₆ And (3) a tower. The distance between each stage of mixer and its corresponding feed location for each stage of reaction was 3 meters.

The heavy-removal tower is a plate-type tower and is used for removing C ₆ The operating temperature of the column was 220℃and the operating pressure 140kPaA; removing C from the ₆ C from the side of the column ₆ The component temperature was 100 ℃.

The total hydrogen and C of the device ₆ The molar ratio of benzene in the components is 1:2.5, the total energy consumption is 180kg, the standard oil/t is hydrogenated terphenyl, the total conversion rate of fresh benzene is 79%, and the total selectivity is 90%.

[ example 2 ]

A1 ten thousand ton/year benzene hydroalkylation hydrogen production terphenyl device adopts the process technology of figure 1, the reactor is 3-level, the three-level reactor is an adiabatic fixed bed reactor, the fresh benzene flow is 1800kg/h, the hydrogen flow of each section is 19kg/h, and the self-stripping C is realized ₆ C extracted from tower ₆ The flow rate of the components is 14000kg/h.

C ₆ The components pass through an interstage cooler step by step, are mixed with 1 st hydrogen in a1 st stage mixer, enter a1 st stage reactor under the conditions of the temperature of 150 ℃ and the pressure of 1.7MPaG, the temperature of the outlet gas of the 1 st stage reactor is 178 ℃, and the outlet gas passes through C ₆ Component coolingThen mixing the mixture with the 2 nd hydrogen in a 2 nd stage mixer, then feeding the mixture into a 2 nd stage reactor under the conditions of the temperature of 150 ℃ and the pressure of 1.65MPaG, wherein the temperature of the outlet gas of the 2 nd stage reactor is 178 ℃, and the outlet gas passes through C ₆ Cooling the components, mixing with 3 rd hydrogen in a 3 rd stage mixer, introducing into a 3 rd stage reactor under the conditions of temperature 151 ℃ and pressure 1.6MPaG, introducing the reaction product into a C removal reactor under the conditions of 179 ℃ and 1.5MPaG ₆ And (3) a tower. The distance between each stage of mixer and its corresponding feed location for each stage of reaction was 2 meters.

The heavy-removal tower is a plate-type tower and is used for removing C ₆ The operating temperature of the column was 200℃and the operating pressure was 130kPaA; removing C from the ₆ C from the side of the column ₆ The temperature of the components is 95 ℃.

The total hydrogen and C of the device ₆ The molar ratio of benzene in the components is 1:4.5, the total energy consumption is 190kg of standard oil/t of hydrogenated terphenyl, the total conversion rate of fresh benzene is 80%, and the total selectivity is 90%.

[ example 3 ]

A1 ten thousand ton/year benzene hydroalkylation hydrogen production terphenyl device adopts the process technology of figure 1, the reactor is five-stage, and the five-stage reactor is an adiabatic fixed bed reactor. Fresh benzene flow rate is 1800kg/h, hydrogen flow rate of each section is 11.5kg/h, and self-stripping C is carried out ₆ C extracted from tower ₆ The flow rate of the components is 4800kg/h.

C ₆ The components pass through an interstage cooler step by step, are mixed with 1 st hydrogen in a1 st stage mixer, enter a1 st stage reactor under the conditions of the temperature of 150 ℃ and the pressure of 1.7MPaG, the temperature of the outlet gas of the 1 st stage reactor is 188 ℃, and the outlet gas passes through C ₆ Cooling the components, mixing with 2 nd hydrogen in a 2 nd stage mixer, introducing into a 2 nd stage reactor under the conditions of 150 ℃ and 1.65MPaG, wherein the temperature of the outlet gas of the 2 nd stage reactor is 188 ℃, and the outlet gas passes through C ₆ Cooling the components, mixing with 3 strands of hydrogen in a 3 rd-stage mixer, introducing into a 3 rd-stage reactor under the conditions of the temperature of 151 ℃ and the pressure of 1.6MPaG, introducing the outlet gas of the 3 rd-stage reactor into the C-stage mixer at the temperature of 188 ℃ and introducing the outlet gas into the C-stage mixer ₆ Cooling the components, mixing with 4 th hydrogen in 4 th stage mixer, and introducing into 1.55MPaG at 152 deg.C4 th stage reactor, outlet gas through C ₆ Cooling the components, mixing with 5 th hydrogen in 5 th stage mixer, introducing into 5 th stage reactor at 152 deg.C under 1.50MPaG, introducing into C-removing reactor at 189 deg.C under 1.5MPaG ₆ And (3) a tower. The distance between each stage of mixer and its corresponding feed location for each stage of reaction was 1 meter.

The heavy-removal tower is a plate-type tower and is used for removing C ₆ The operating temperature of the column was 180℃and the operating pressure was 80kPaA; removing C from the ₆ C from the side of the column ₆ The component temperature was 80 ℃.

The total hydrogen and C of the device ₆ The molar ratio of benzene in the components is 1:2.5, the total energy consumption is 175kg of standard oil/t of hydrogenated terphenyl, the total conversion rate of fresh benzene is 81%, and the total selectivity is 91%.

[ example 4 ]

A1 ten thousand ton/year benzene hydroalkylation hydrogen production terphenyl device adopts the process technology of figure 1, the reactor is five-stage, and the five-stage reactor is an adiabatic fixed bed reactor. Fresh benzene flow rate is 1800kg/h, hydrogen flow rate of each section is 11.5kg/h, and self-stripping C is carried out ₆ C extracted from tower ₆ The flow rate of the components is 4800kg/h.

C ₆ The components pass through an interstage cooler step by step, are mixed with 1 st hydrogen in a1 st stage mixer, enter a1 st stage reactor under the conditions of the temperature of 180 ℃ and the pressure of 2MPaG, the temperature of the outlet gas of the 1 st stage reactor is 219 ℃, and the outlet gas passes through C ₆ Cooling the components, mixing with 2 nd hydrogen in a 2 nd stage mixer, introducing into a 2 nd stage reactor under the conditions of temperature 181 ℃ and pressure 1.95MPaG, wherein the temperature of the outlet gas of the 2 nd stage reactor is 219 ℃, and the outlet gas passes through C ₆ Cooling the components, mixing with 3 strands of hydrogen in a 3 rd-stage mixer, introducing into a 3 rd-stage reactor under the conditions of temperature 181 ℃ and pressure 1.90MPaG, introducing the outlet gas of the 3 rd-stage reactor into a C-stage mixer at the temperature of 220 DEG, introducing the outlet gas into a C-stage mixer ₆ Cooling the components, mixing with 4 th hydrogen in 4 th stage mixer, introducing into 4 th stage reactor under the condition of 182 deg.C and 1.85MPaG pressure, and introducing the outlet gas into C ₆ Cooling the components, mixing with 5 th hydrogen in 5 th stage mixer, and under the condition of temperature 183 deg.C and pressure 1.80MPaGDown into a 5 th stage reactor, and the reaction product enters into a C removal reactor at 221 ℃ and 1.75MPaG ₆ And (3) a tower. The distance between each stage of mixer and its corresponding feed location for each stage of reaction was 1 meter.

The heavy-removal tower is a plate-type tower and is used for removing C ₆ The operating temperature of the column was 180℃and the operating pressure was 80kPaA; removing C from the ₆ C from the side of the column ₆ The component temperature was 80 ℃.

The total hydrogen and C of the device ₆ The molar ratio of benzene in the components is 1:1.5, the total energy consumption is 170kg of standard oil/t of hydrogenated terphenyl, the total conversion rate of fresh benzene is 80%, and the total selectivity is 90%.

Comparative example 1

The procedure of example 2 was repeated, except that: fresh benzene has not undergone the process of removing C ₆ The column is associated with the removal of C ₆ C from the side of the column ₆ The components are mixed prior to entering cooler E102. Other conditions were unchanged.

The reactor is unstable in operation, the temperature of the catalyst bed is not uniform, the catalyst is deactivated due to moisture brought by fresh benzene after one month of operation, and the device is stopped for changing the catalyst.

Comparative example 2

The procedure of example 2 was repeated, except that: de-C ₆ The side offtake position of the column is in the same horizontal direction (i.e. the same height) as the feed position of fresh benzene. Other conditions were unchanged.

De-C ₆ Column operation is unstable, feed trays often experience evacuation and the production of C ₆ The components still contain 500-1000ppm of water, the reactor is unstable to operate, the temperature of the catalyst bed is not uniform, the catalyst is deactivated by the water brought by fresh benzene after three months of operation, and the device is stopped for changing the catalyst.

The invention has been described in detail in connection with the specific embodiments and exemplary examples thereof, but such description is not to be construed as limiting the invention. It will be understood by those skilled in the art that various equivalent substitutions, modifications or improvements may be made to the technical solution of the present invention and its embodiments without departing from the spirit and scope of the present invention, and these fall within the scope of the present invention. The scope of the invention is defined by the appended claims.

Claims (15)

1. A method of preparing hydrogenated terphenyl comprising: with hydrogen and from dec ₆ Column C ₆ The components are raw materials for N-stage reaction, N is more than or equal to 2, and the reaction product after the N-stage reaction is finished enters the C removal device ₆ A tower; wherein, at the dec ₆ In the tower, fresh benzene is taken as a feed, and the C is laterally extracted ₆ The components are extracted from the tower top materials, and the product containing hydrogenated terphenyl is extracted from the tower bottom.

2. The process of claim 1, comprising N-stage reactions, the hydrogen being fed separately to each stage reaction in N separate streams from the dec ₆ Column C ₆ The components enter each section in turn for reaction in one strand; preferably, n=2 to 6, preferably 2 to 4.

3. The method according to claim 1, wherein the reaction product of the previous reaction is subjected to a cooling treatment before the reaction product of the previous reaction enters the next reaction; preferably, to come from dec ₆ Said C of the column ₆ The components are cooling medium; more preferably, in said C ₆ The components are subjected to heat exchange treatment with the reaction products of the (N-1) th section reaction, the (N-2) th section reaction, … and the 1 st section reaction in sequence before entering each section reaction, and then enter each section reaction in sequence after heat exchange.

4. The method of claim 1, wherein the step of determining the position of the substrate comprises,

the total amount of the hydrogen and C ₆ The mole ratio of benzene in the component is 1 (1.1-6), preferably 1 (1.5-4.5); and/or the number of the groups of groups,

the hydrogen feed amount in each reaction is the same or different, preferably the same.

5. The method of claim 1, wherein hydrogen is reacted with the C prior to entering each stage of reaction ₆ Component or hydrogen and upper partThe reaction products of the first-stage reaction are firstly mixed in a mixer; preferably, the distance between each stage of mixer and its corresponding feed position for each stage of reaction is greater than or equal to 0.5 meter, preferably greater than or equal to 1 meter.

6. The process according to claim 1, wherein the reaction inlet temperature of each stage reaction is independently 130 to 280 ℃, preferably 150 to 220 ℃; the pressure of each stage reaction is independently 0.5 to 2.5MPaG, preferably 1.0 to 2.0MPaG.

7. The method of claim 1, wherein the step of determining the position of the substrate comprises,

the said dec ₆ The operating temperature of the column is 150 to 250 ℃, preferably 180 to 220 ℃; the operating pressure is 20 to 200kPaA, preferably 80 to 140kPaA; and/or the number of the groups of groups,

removing C from the ₆ C from the side of the column ₆ The temperature of the components is 60-120 ℃.

8. The method according to any one of claims 1 to 7, wherein, in said dec ₆ The reaction product feed location for the N-stage reaction is below the feed location for fresh benzene in the column; preferably, the reaction product of the N-stage reaction is removed from the C ₆ The lower part of the tower is introduced, and the fresh benzene is removed from C ₆ The upper part of the tower enters; more preferably, in said dec ₆ In the column, the fresh benzene is fed at a position C ₆ Below the side offtake position of the components.

9. The process of claim 8, wherein the N-stage reactions are each independently carried out in the presence of a catalyst; preferably, the catalyst is a metal supported bi-functional solid acid hydroalkylation catalyst;

more preferably, the catalyst comprises a support, a binder and an active component, wherein the support is a silicone microporous zeolite, the active component comprises a combination of a noble metal active component (preferably selected from at least one of palladium, ruthenium or platinum) and a non-noble metal (preferably selected from nickel, copper, cobalt), the binder is selected from at least one of zinc oxide, aluminum oxide or titanium oxide;

most preferably, the carrier is present in an amount of 30 to 90wt%, the binder is present in an amount of 8 to 60wt%, the noble metal active component is present in an amount of 0.05 to 5wt%, and the non-noble metal active component is present in an amount of 10 to 60wt%, based on 100wt% of the total weight of the catalyst.

10. A system for preparing hydrogenated terphenyl, preferably for carrying out the process of any one of claims 1 to 9, comprising N-stage series reactors and a dec ₆ The tower, wherein, along the material flow direction, is 1 st level reactor, …, (N-1) th level reactor and N level reactor in proper order, and N is greater than or equal to 2.

11. The system of claim 10, wherein the system further comprises a controller configured to control the controller,

the N-stage series reactors are N reactors connected in series, wherein N=2-6, preferably 2-4; and/or the number of the groups of groups,

a feed inlet is arranged at the upper part or the top of each stage of reactor, and a discharge outlet is arranged at the lower part or the bottom of each stage of reactor; preferably, the discharge port of the upper stage reactor is connected with the feed port of the lower stage reactor.

12. The system of claim 10, wherein, in said dec-ating ₆ The tower is provided with a reaction material inlet, a fresh benzene inlet, a side line outlet, an outer tower top outlet and an outer tower kettle outlet;

preferably, in said dec ₆ The reaction material inlet is arranged below the fresh benzene inlet, and the side-line extraction outlet is arranged above the fresh benzene inlet;

more preferably, in said dec ₆ On the tower, the reaction material inlet is arranged on the C-removing device ₆ The lower part of the tower, the fresh benzene inlet and the side-line extraction outlet are arranged at the C-removing part ₆ The upper part of the tower and the side-draw outlet are arranged above the fresh benzene inlet.

13. The system of claim 12, wherein the system further comprises a controller configured to control the controller,

De-C ₆ The reaction material inlet of the tower is connected with the discharge port of the Nth stage reactor; and/or take off C ₆ The tower top outer extraction port of the tower is arranged at the tower top; and/or take off C ₆ The tower kettle outer discharge and extraction outlet of the tower is arranged at the tower kettle; and/or the number of the groups of groups,

at least one cooler is arranged between two adjacent stages of reactors independently; preferably, let-off C ₆ The side offtake of the tower, all coolers and the feed inlet of the 1 st stage reactor are connected in sequence through pipelines.

14. The system according to any of claims 10 to 13, characterized in that a mixer is provided separately before the feed inlet of each stage reactor; preferably, the distance between each mixer and the feed inlet of its corresponding reactor of each stage is greater than or equal to 0.5 meter, preferably greater than or equal to 1 meter.

15. Hydrogenated terphenyl obtainable by the process according to any one of claims 1 to 9 or by the system according to any one of claims 10 to 14.

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