A method, system and resulting hydrogenated terphenyls for making hydrogenated terphenyls

Abstract

The application discloses a method and system for preparing hydrogenated terphenyl and the obtained hydrogenated terphenyl, and the method comprises the following steps: carrying out N-stage reactions by taking hydrogen and C6 components from a C6 removal tower as raw materials, wherein N is greater than or equal to 2, and the reaction product enters the C6 removal tower after the N-stage reactions are completed; wherein, fresh benzene is taken as the feed in the C6 removal tower, the C6 components are taken out from a side line, overhead materials are taken out, and a product containing hydrogenated terphenyl is taken out from a tower bottom. The application can realize the preparation of the hydrogenated terphenyl by taking cheap benzene and hydrogen as raw materials under the conditions that the temperature is 130-280 DEG C (preferably 150-220 DEG C) and the pressure is 0.5-2.5 MPaG (preferably 1.0-2.0 MPaG), thereby avoiding the high-temperature cracking operation of benzene.

Description

Technical Field

[0001] This invention relates to the preparation of hydrogenated terphenyl, and more particularly to a method, system for preparing hydrogenated terphenyl and the resulting hydrogenated terphenyl. Background Technology

[0002] Heat transfer oil, as an effective heat transfer medium, is widely used in industries such as petroleum, petrochemicals, manufacturing, and food. Currently, domestic heat transfer oils are mainly mineral oil-based, with an operating temperature below 300℃. However, with increasingly stringent environmental requirements in China and the continuous development of new industries and technologies such as photovoltaic power generation, energy storage, and polyester, the demand for synthetic high-temperature heat transfer oils with long service life and higher operating temperatures (≥330℃), such as L-QD340, has increased significantly. Hydrogenated terphenyl is the main base oil raw material for L-QD340, and currently, this product mainly relies on imports, with domestic production capacity being extremely limited and unable to meet the current urgent market demand.

[0003] Patent CN109053355A relates to a method for continuous distillation purification of biphenyl. The method involves passing a feed liquid through a benzene stripping tower, a biphenyl tower, and a terphenyl tower connected in series. The feed liquid is added from the middle of the benzene stripping tower. A small amount of hydrogen, benzene, and biphenyl are collected from the top of the stripping tower. The remaining biphenyl, terphenyl, and terphenyl mixture, along with sand tar, are collected from the bottom of the tower and fed into the biphenyl tower. A small amount of benzene-containing biphenyl mixture is collected from the top of the biphenyl tower and returned to the reaction liquid storage tank. High-purity biphenyl is collected from the side stream. A mixture of terphenyl and sand tar is collected from the bottom of the tower and fed into the terphenyl tower. A small amount of terphenyl-containing biphenyl mixture is collected from the top of the terphenyl tower and returned to the biphenyl tower feed inlet. A high-purity terphenyl mixture is collected from the side stream. Heavy sand tar is collected from the bottom of the tower. By setting up side-stream extraction, high-purity biphenyl products with a mass fraction of over 99.9% can be obtained while being suitable for large-scale production and having high operational flexibility. Compared with traditional batch processes, its unit product energy consumption can be reduced by 30-40%, which has extremely high economic value.

[0004] Patent CN103804114A describes a process where benzene is cracked at high temperature to obtain reaction products containing benzene, biphenyl, and terphenyl. These products are then cooled and distilled to obtain terphenyl, which is subsequently hydrogenated in a high-pressure hydrogenation reactor to obtain hydrogenated terphenyl. This patent utilizes a high-performance, energy-saving tubular reactor, which reduces energy consumption to some extent. While all the above methods can yield hydrogenated terphenyl, the raw material still originates from the high-temperature cracking of benzene.

[0005] However, in the existing technology, the traditional industrial production method of hydrogenated terphenyl is the high-temperature cracking hydrogenation of benzene, which is a by-product of biphenyl production. This method has a high cracking temperature (≥800℃), high energy consumption, and low overall yield (4%). Summary of the Invention

[0006] To overcome the problems existing in the prior art, this invention provides a method, system, and the resulting hydrogenated terphenyl, mainly addressing the issues of high raw material costs, high energy consumption, and low overall yield in previous technologies. This invention employs an N-stage reactor series reaction. The feed to the first-stage reactor is hydrogen and C6 components from a C6 removal tower. Subsequent reactors are fed hydrogen and reaction products from the previous stage reactor. The reactors have interstage coolers, with the C6 components from the C6 removal tower serving as the cooling medium. The reaction products from the Nth-stage reactor and fresh benzene enter the C6 removal tower from different locations. The separated C6 components are returned to the first-stage reactor via the coolers, and the bottom stream is fed into subsequent processes for the recovery of hydrogenated terphenyl. This technical solution effectively solves the problem and can be used in the industrial production of hydrogenated terphenyl.

[0007] One objective of this invention is to provide a method for preparing hydrogenated terphenyl, comprising: conducting an N-stage reaction using hydrogen and C6 components from a C6 removal tower as raw materials, where N ≥ 2; and the reaction product after the Nth stage reaction (i.e., the last stage) is introduced into the C6 removal tower; wherein, in the C6 removal tower, fresh benzene is used as feed, the C6 components are collected from the side stream, the top material is collected externally, and the product containing hydrogenated terphenyl is collected from the bottom of the tower.

[0008] The reaction proceeds sequentially along the flow path of hydrogen and C6 components: stage 1, ..., stage (N-1), and stage N. In the C6 removal tower, the heavy components in the N-stage reaction product (including components above C6 but without C6, such as C...) 12 and C 18 The product is collected from the bottom of the column, while components with C6 and above (including unreacted benzene and cyclohexane) are in the upper part of the C6 removal column. Therefore, the C6 component contains benzene and cyclohexane, wherein the benzene is partly fresh benzene and partly unreacted benzene and byproduct cyclohexane from the Nth stage reaction products (the cyclohexane is recycled back to subsequent reactions and has no effect on the reaction).

[0009] In this invention, hydrogen is reacted directly with benzene to form hydrogenated terphenyl, thus avoiding the use of high-temperature pyrolysis.

[0010] In a preferred embodiment, the method includes N stages of reaction, with hydrogen entering each stage of reaction independently in N separate streams, and C6 components from the C6 removal tower entering each stage of reaction sequentially in one stream; that is, the feed for the first stage of reaction is hydrogen and the C6 components, and the feed for each subsequent stage of reaction is hydrogen and the reaction products from the previous stage of reaction.

[0011] In a further preferred embodiment, N = 2 to 6, preferably 2 to 4, for example N = 2, 3, 4, 5 or 6.

[0012] In a preferred embodiment, the reaction product of the previous reaction is cooled before it enters the next reaction.

[0013] In this invention, the reaction between hydrogen and benzene is an exothermic reaction. After each stage of the reaction, the temperature of the reaction product is high. If the temperature is not lowered and the next stage of the reaction is directly started, the temperature of the subsequent reaction will continue to rise, which will affect the performance of the catalyst.

[0014] In a further preferred embodiment, the C6 component from the C6 removal tower is used as the cooling medium.

[0015] In a further preferred embodiment, the C6 component is subjected to heat exchange treatment sequentially with the reaction products of the (N-1)th reaction, the (N-2)th reaction, ..., the first reaction before entering each reaction stage, and then enters each reaction stage in one stream.

[0016] In this process, the C6 component needs to be heated to near the reaction temperature before entering the first stage reaction. Using existing methods, this might involve heating the C6 component with an electric heater before it enters the first stage reaction. However, in this application, heat exchange is performed between the C6 component and the reaction products between stages. This allows the C6 component to be heated to the required temperature while simultaneously cooling the reaction products, resulting in significant energy savings.

[0017] In a preferred embodiment, the N-stage reactions each proceed independently in the presence of a catalyst.

[0018] In a further preferred embodiment, the catalyst is a metal-supported bifunctional solid acid hydrogenation alkylation catalyst.

[0019] In this invention, the metal-supported bifunctional solid acid hydrogenation alkylation catalyst is selected from those disclosed in the prior art, preferably but not limited to: the catalyst includes a support, a binder and an active component, wherein the support is an organosilicon microporous zeolite, the active component includes 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 or cobalt), and the binder is selected from at least one of zinc oxide, aluminum oxide or titanium oxide.

[0020] Preferably, based on 100 wt% of the total weight of the catalyst, the content of the support is 30-90 wt%, the content of the binder is 8-60 wt%, the content of the noble metal active component is 0.05-5 wt%, and the content of the non-noble metal active component is 10-60 wt%.

[0021] For example, based on 100 wt% of the total weight of the catalyst, the content of the support is 30 wt%, 40 wt%, 50 wt%, 60 wt%, 70 wt%, 80 wt%, or 90 wt%, the content of the binder is 8 wt%, 10 wt%, 20 wt%, 30 wt%, 40 wt%, 50 wt%, or 60 wt%, the content of the noble metal active component is 0.05 wt%, 0.1 wt%, 0.5 wt%, 1 wt%, 2 wt%, 3 wt%, 4 wt%, or 5 wt%, and the content of the non-noble metal active component is 10 wt%, 25 wt%, 30 wt%, 35 wt%, 40 wt%, 45 wt%, 50 wt%, 55 wt%, or 60 wt%.

[0022] In this invention, the organosilicon microporous zeolite can be purchased directly or prepared using methods disclosed in the prior art. Preferably, the organosilicon microporous zeolite has the following molar composition: (1 / n)Al₂O₃:SiO₂:(m / n)R, where n = 5–250, m = 0.01–50, and R is at least one of an alkyl or phenyl group having 1–8 carbon atoms; the Si of the organosilicon microporous zeolite... 29 The NMR solid-state NMR spectrum contains at least one Si in the range of -80 to +50 ppm. 29 Nuclear magnetic resonance peaks; the X-ray diffraction pattern of the organosilicon microporous zeolite has d-spacing maximum values ​​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 Å.

[0023] In this invention, preferably, the catalyst is prepared as follows: First, the above-mentioned organosilicon microporous zeolite is ground into powder, and a mixed solution containing noble metal active components and non-noble metal active components is prepared. The prepared mixed solution is sprayed onto the organosilicon microporous zeolite powder, and the organosilicon microporous zeolite is continuously stirred during the spraying process. After drying at room temperature and pressure for 1-20 hours, it is then dried at 100-150°C for 1-20 hours and ground into powder. This powder is mixed with a binder, kneaded into shape, dried, and then calcined in air at 400-550°C for 1-10 hours. After exchange in an ammonium salt solution, it is washed, dried, and then calcined in air at 480°C for 5 hours to obtain the finished catalyst.

[0024] In a preferred embodiment, the N-stage reactions are carried out independently in a reactor. Preferably, the reactor is an adiabatic fixed-bed reactor. More preferably, each reactor is independently filled with the catalyst.

[0025] In a preferred embodiment, the total amount of hydrogen used is in a molar ratio of 1:(1.1 to 6) to benzene in the C6 component, preferably 1:(1.5 to 4.5).

[0026] For example, the molar ratio of hydrogen to benzene in component C6 is 1:1.1, 1:1.2, 1:1.5, 1:2, 1:2.5, 1:3, 1:3.5, 1:4, 1:4.5, 1:5, 1:5.5 or 1:6.

[0027] If the total amount of hydrogen used is too high compared to the molar ratio of benzene in the C6 component, the energy consumption will be high; if it is too low, the catalyst life will be greatly reduced.

[0028] In a further preferred embodiment, the amount of hydrogen fed in each stage of the reaction may be the same or different, preferably the same.

[0029] In a preferred embodiment, the hydrogen gas is mixed with the C6 component or the hydrogen gas with the reaction products of the previous reaction before entering each stage of the reaction.

[0030] In the first stage of the reaction, hydrogen gas and (after heat exchange) C6 component are first mixed in the mixer; before the second to Nth stages of the reaction, hydrogen gas and (after heat exchange) the reaction products of the previous stage are first mixed in the mixer.

[0031] In a further preferred embodiment, the distance between each stage mixer and the feed position of its corresponding reaction stage is greater than or equal to 0.5 meters, preferably greater than or equal to 1 meter.

[0032] This allows for the control of uniform material mixing, especially the uniform mixing of hydrogen and benzene.

[0033] In a preferred embodiment, the reaction inlet temperature of each reaction stage is independently 130–280°C, and the pressure of each reaction stage is independently 0.5–2.5 MPaG.

[0034] In a further preferred embodiment, the reaction inlet temperature of each reaction stage is independently 150–220°C, and the pressure of each reaction stage is independently 1.0–2.0 MPaG.

[0035] In a further preferred embodiment, the reaction inlet temperature of each reaction stage is independently 150–180°C, and the pressure of each reaction stage is independently 1.2–1.7 MPaG.

[0036] For example, the reaction inlet temperature of each stage of the reaction is independently 150°C, 160°C, 170°C, 180°C, 200°C or 220°C, and the pressure of each stage of the reaction is independently 1.0 MPaG, 1.2 MPaG, 1.4 MPaG, 1.6 MPaG, 1.8 MPaG or 2.0 MPaG.

[0037] In a preferred embodiment, in the C6 removal tower, the feed position of the N-stage reaction product is below the feed position of the fresh benzene.

[0038] In a further preferred embodiment, the reaction products of the N-stage reaction enter from the lower part of the C6 removal tower, and the fresh benzene enters from the upper part of the C6 removal tower.

[0039] The N-stage reaction may produce unreacted benzene, the target product hydrogenated terphenyl, byproduct cyclohexane, and byproduct C. 12 Components and C except for hydrogenated terphenyl 18 Components. Through extensive experimentation, the inventors discovered that when the reaction products of the N-stage reaction are fed into the lower part of the C6 removal tower and the tower is controlled within a suitable temperature and pressure range, unreacted benzene and byproduct cyclohexane can enter the upper part of the C6 removal tower and be collected as a side stream along with fresh benzene, while the target product, hydrogenated terphenyl, and byproduct C... 12 Components and C except for hydrogenated terphenyl 18 The components are collected from the bottom of the column. In this way, not only are unreacted benzene and by-product cyclohexane separated from the product, but the unreacted benzene can also be further recycled and utilized.

[0040] In a further preferred embodiment, in the C6 removal tower, the feed location of fresh benzene is below the side-stream exit location of the C6 component.

[0041] In this process, fresh benzene rises from the feed location to the side feed outlet and is collected via a side stream. Fresh benzene contains 1000–2000 ppm of water (which significantly affects my catalyst) and other light component impurities. After treatment in the C6 removal tower, the water and light component impurities in the fresh benzene are discharged from the top of the tower, meaning the C6 removal tower purifies the fresh benzene. Therefore, the C6 component collected via the side stream does not contain water or other light components.

[0042] In a preferred embodiment, the operating temperature of the C6 removal tower is 150–250°C, and the operating pressure is 20–200 kPaA.

[0043] In a further preferred embodiment, the operating temperature of the C6 removal tower is 180–220°C, and the operating pressure is 80–140 kPaA.

[0044] For example, the operating temperature of the C6 removal tower is 150°C, 180°C, 200°C, 220°C, 240°C or 250°C, and the operating pressure is 20 kPaA, 50 kPaA, 80 kPaA, 100 kPaA, 120 kPaA, 150 kPaA, 180 kPaA or 200 kPaA.

[0045] In this invention, the deweight removal tower is preferably a plate tower.

[0046] In a preferred embodiment, the temperature of the C6 component extracted from the C6 removal tower side stream is 60–120°C, for example, 60°C, 70°C, 80°C, 90°C, 100°C, 110°C, or 120°C.

[0047] In this invention, the material extracted from the bottom of the C6 removal tower enters the subsequent process for recovering hydrogenated terphenyl (using any method disclosed in the prior art). Preferably, the subsequent process includes vacuum distillation of hydrogenated terphenyl and heat recovery.

[0048] In a preferred embodiment, the method is carried out as follows: N series reactions are employed, where N is an integer greater than or equal to 2; the reaction process includes the following steps:

[0049] 1) After hydrogen and C6 components are mixed, they come into contact with the catalyst and enter the first stage reaction to generate the first-stage stream;

[0050] 2) After the first-stage stream is cooled by the C6 component, it is mixed with hydrogen and comes into contact with the catalyst to enter the second stage reaction, generating the second-stage stream;

[0051] 3) When N equals 2, the second-stage stream enters step 4). When N is greater than 2, the second-stage stream is cooled by C6 component, mixed with hydrogen, and contacted with catalyst to enter the third stage reaction to generate the third-stage stream. Step 3) is repeated until the reacted material enters the Nth stage reaction to generate the Nth-stage stream.

[0052] 4) The Nth stage stream and fresh benzene enter the C6 removal tower at different locations;

[0053] 5) The C6 component is taken out from the side stream of the C6 removal tower, and after being heat-exchanged with the reaction products of each stage, it is returned to the first stage reaction. The top material of the C6 removal tower is extracted, and the bottom material enters the subsequent section to recover hydrogenated terphenyl.

[0054] A second objective of this invention is to provide a system for preparing hydrogenated terphenyl, preferably for carrying out the method described in one objective of this invention. The system includes N-stage reactors in series and a C6 removal tower, wherein the reactors are arranged sequentially along the material flow direction as stage 1, ..., stage (N-1) and stage N, where N ≥ 2.

[0055] In a preferred embodiment, the N-stage series reactor is N reactors connected in series, where N = 2 to 6, preferably 2 to 4, for example N = 2, 3, 4, 5 or 6.

[0056] In a preferred embodiment, a feed inlet is provided at the top or upper part of each reactor stage, and a discharge outlet is provided at the bottom or lower part of each reactor stage.

[0057] In a further preferred embodiment, the outlet of the previous stage reactor is connected to the inlet of the next stage reactor.

[0058] In a preferred embodiment, the C6 removal tower is provided with a reactant inlet, a fresh benzene inlet, a side stream outlet, a top outlet, and a bottom outlet.

[0059] In a further preferred embodiment, on the C6 removal tower, the reactant inlet is located below the fresh benzene inlet, and the side stream outlet is located above the fresh benzene inlet.

[0060] In a further preferred embodiment, in the C6 removal tower, the reactant inlet is located at the lower part of the C6 removal tower, the fresh benzene inlet and the side stream outlet are located at the upper part of the C6 removal tower, and the side stream outlet is located above the fresh benzene inlet.

[0061] In a preferred embodiment, the reactant inlet of the C6 removal column is connected to the outlet of the Nth stage reactor; and / or, the top outlet of the C6 removal column is located at the top of the column; and / or, the bottom outlet of the C6 removal column is located at the bottom of the column.

[0062] In a preferred embodiment, at least one cooler is independently provided between each of two adjacent reactor stages, i.e., at least (N-1) coolers are provided in total.

[0063] In a further preferred embodiment, the side outlet of the C6 removal tower, all coolers, and the feed inlet of the first-stage reactor are sequentially connected by pipelines; preferably, in the direction opposite to the material flow between each stage reactor, the side outlet of the C6 removal tower is sequentially connected to each cooler and then to the feed inlet of the first-stage reactor.

[0064] In this way, the material extracted from the side outlet passes through each cooler in the opposite direction to the material flow between each stage reactor and finally enters the first stage reactor.

[0065] In a preferred embodiment, a mixer is independently provided before the feed inlet of each stage reactor for mixing hydrogen with the C6 component or the reaction products of the previous stage reactor.

[0066] In a further preferred embodiment, the distance between each mixer and the feed inlet of its corresponding stage reactor is greater than or equal to 0.5 meters, preferably greater than or equal to 1 meter.

[0067] A third objective of this invention is to provide hydrogenated terphenyl obtained using the method or system described in this invention.

[0068] The endpoints and any values ​​of the ranges disclosed in this invention are not limited to the precise ranges or values; these ranges or values ​​should be understood to include values ​​close to these ranges or values. For numerical ranges, the endpoint values ​​of the various ranges, the endpoint values ​​of the various ranges and individual point values, and individual point values ​​can be combined with each other to obtain one or more new numerical ranges, which should be considered as specifically disclosed herein. In the following, various technical solutions can, in principle, be combined with each other to obtain new technical solutions, which should also be considered as specifically disclosed herein.

[0069] Compared with the prior art, the present invention has the following beneficial effects:

[0070] (1) Using inexpensive benzene and hydrogen as raw materials, hydrogenated terphenyl can be prepared at a temperature of 130-280℃ (preferably 150-220℃) and a pressure of 0.5-2.5MPaG (preferably 1.0-2.0MPaG), avoiding the high-temperature cracking operation of benzene;

[0071] (2) Using C6 component as cooling medium to achieve precise control of reaction temperature and effectively utilize reaction heat;

[0072] (3) The total conversion rate of benzene is higher than 78%, and the total yield of hydrogenated terphenyl is higher than 70%;

[0073] (4) Compared with traditional methods, this method has the advantages of cheap and readily available raw materials, mild reaction conditions, easy control of heat removal, low energy consumption and high selectivity of target products. Attached Figure Description

[0074] Figure 1 A schematic diagram of one embodiment of the system described in this invention is shown. Attached image description:

[0076] R-101 - Stage 1 reactor, R-102 - Stage 2 reactor, R-103 - Stage 3 reactor, M - C6 removal tower, E-101 - Stage 1 cooler, E-102 - Stage 2 cooler, 1 - Hydrogen, 2 - Reaction product of Stage 1 reaction, 3 - Reaction product of Stage 2 reaction, 4 - Reaction product of Stage 3 reaction, 5 - Fresh benzene, 6 - C6 component, 7 - Top product of tower, 8 - Bottom product of tower. Detailed Implementation

[0077] The present invention will now be described in detail with reference to specific embodiments. It should be noted that the following embodiments are only used to further illustrate the present invention and should not be construed as limiting the scope of protection of the present invention. Some non-essential improvements and adjustments made by those skilled in the art based on the content of the present invention are still within the scope of protection of the present invention.

[0078] It should also be noted that the various specific technical features described in the following embodiments can be combined in any suitable manner without contradiction. To avoid unnecessary repetition, the various possible combinations will not be described separately in this invention.

[0079] Furthermore, various embodiments of the present invention can be combined in any way, as long as they do not violate the spirit of the present invention. The resulting technical solutions are part of the original disclosure of this specification and also fall within the protection scope of the present invention.

[0080] The catalysts used in the examples and comparative examples were prepared as follows:

[0081] (1) Synthesis of organosilicon microporous zeolite: 3.0 g of alumina was dissolved in 450 g of water, and 16.0 g of sodium hydroxide was added to dissolve it. Then, under stirring, 34.7 g of hexamethyleneimine, 60 g of solid silicon dioxide, and 5.9 g of dimethyldiethoxysilane were added. The molar ratio of the reactants was: SiO2 / Al2O3 = 30, NaOH / SiO2 = 0.2, dimethyldiethoxysilane / SiO2 = 0.04, hexamethyleneimine / SiO2 = 0.35, and H2O / SiO2 = 25. After the reaction mixture was stirred evenly, it was transferred to a stainless steel reactor and crystallized at 145 °C for 70 hours under stirring. After removal, it was filtered, washed, and dried. Chemical analysis showed that the molar ratio of SiO2 / Al2O3 was 30.1.

[0082] (2) Synthesis of the catalyst: Prepare a 40 ml mixed solution of ruthenium chloride and copper chloride, with Ru content of 0.25 g, CuCl2 content of 9.95 g, and a Ru to Cu molar ratio of 1:30. Take 32.50 g of organosilicon microporous zeolite powder and spray the prepared mixed solution onto the organosilicon microporous zeolite powder, stirring the organosilicon microporous zeolite continuously during the spraying process. After drying at room temperature and pressure for 10 hours, dry at 120℃ for 10 hours and grind into powder. Add 8.98 g of alumina and mix, then add dilute nitric acid solution to knead and extrude into strips of φ1.6×2 mm. After drying, calcine at 550℃ for 5 hours, then exchange with 1M ammonium nitrate 5 times, filter, and dry. Dry at 120℃ for 12 hours and calcine at 480℃ for 6 hours to prepare the required catalyst. The composition, by weight percentage, is 0.52% Ru, 12.5% ​​Cu, 18.7% binder, 67.7% organosilicon microporous zeolite, and other impurities.

[0083]

Example 1

[0084] A 10,000-ton / year benzene hydrogenation alkylation to terphenyl hydrogenation unit adopts... Figure 1 The process technology is advanced, but the reactor is a four-stage reactor, all of which are adiabatic fixed-bed reactors. The fresh benzene flow rate is 1800 kg / h, the hydrogen flow rate of each stage is 14 kg / h, and the C6 component flow rate from the C6 desorption tower is 8000 kg / h.

[0085] The C6 component passes through interstage coolers and is mixed with the first hydrogen stream in the first-stage mixer. It then enters the first-stage reactor at 150°C and 1.7 MPaG, with an outlet gas temperature of 184°C. After being cooled by the C6 component, the outlet gas is mixed with the second hydrogen stream in the second-stage mixer and enters the second-stage reactor at 150°C and 1.65 MPaG, with an outlet gas temperature of 187°C. After being cooled by the C6 component, the gas is mixed with the third stream of hydrogen in the third-stage mixer. It then enters the third-stage reactor at a temperature of 151°C and a pressure of 1.6 MPaG. The outlet gas temperature of the third-stage reactor is 188°C. After being cooled by the C6 component, the outlet gas is mixed with the fourth stream of hydrogen in the fourth-stage mixer. It then enters the fourth-stage reactor at a temperature of 152°C and a pressure of 1.55 MPaG. The reaction products enter the C6 removal tower at 188°C and 1.5 MPaG. The distance between each stage mixer and its corresponding feed location for each reaction stage is 3 meters.

[0086] The deweight removal tower is a plate tower, and the operating temperature of the C6 removal tower is 220℃ and the operating pressure is 140kPaA; the temperature of the C6 component sampled from the side line of the C6 removal tower is 100℃.

[0087] The unit has a total hydrogen to benzene molar ratio of 1:2.5, a total energy consumption of 180 kg standard oil / t hydrogenated terphenyl, a total fresh benzene conversion rate of 79%, and a total selectivity of 90%.

[0088]

Example 2

[0089] A 10,000-ton / year benzene hydrogenation alkylation to terphenyl hydrogenation unit adopts... Figure 1 The process technology involves a three-stage reactor, all of which are adiabatic fixed-bed reactors. The fresh benzene flow rate is 1800 kg / h, the hydrogen flow rate for each stage is 19 kg / h, and the C6 component flow rate from the C6 desorption tower is 14000 kg / h.

[0090] The C6 component passes through interstage coolers and is mixed with the first hydrogen stream in the first-stage mixer. It then enters the first-stage reactor at 150°C and 1.7 MPaG. The outlet gas temperature of the first-stage reactor is 178°C. After being cooled by the C6 component, the outlet gas is mixed with the second hydrogen stream in the second-stage mixer and enters the second-stage reactor at 150°C and 1.65 MPaG. The outlet gas temperature of the second-stage reactor is 178°C. After being cooled by the C6 component, the outlet gas is mixed with the third hydrogen stream in the third-stage mixer and enters the third-stage reactor at 151°C and 1.6 MPaG. The reaction products enter the C6 removal tower at 179°C and 1.5 MPaG. The distance between each mixer and its corresponding feed location for each reaction stage is 2 meters.

[0091] The deweight removal tower is a plate tower, and the operating temperature of the C6 removal tower is 200℃ and the operating pressure is 130kPaA; the temperature of the C6 component sampled from the side line of the C6 removal tower is 95℃.

[0092] The unit has a total hydrogen molar ratio of 1:4.5 to benzene in the C6 component, a total energy consumption of 190 kg standard oil / t hydrogenated terphenyl, a total fresh benzene conversion rate of 80%, and a total selectivity of 90%.

[0093]

Example 3

[0094] A 10,000-ton / year benzene hydrogenation alkylation to terphenyl hydrogenation unit adopts... Figure 1 The process technology involves a five-stage reactor, all of which are adiabatic fixed-bed reactors. The fresh benzene flow rate is 1800 kg / h, the hydrogen flow rate for each stage is 11.5 kg / h, and the C6 component flow rate from the C6 removal tower is 4800 kg / h.

[0095] The C6 component passes through interstage coolers, mixes with the first hydrogen stream in the first-stage mixer, and then enters the first-stage reactor at 150°C and 1.7 MPaG. The outlet gas temperature of the first-stage reactor is 188°C. After being cooled by the C6 component, the outlet gas is mixed with the second hydrogen stream in the second-stage mixer and then enters the second-stage reactor at 150°C and 1.65 MPaG. The outlet gas temperature of the second-stage reactor is 188°C. After being cooled by the C6 component, the outlet gas is mixed with three hydrogen streams in the third-stage mixer. The gas enters the third-stage reactor at 151°C and 1.6 MPaG. The outlet gas temperature of the third-stage reactor is 188°C. After being cooled by the C6 component, the outlet gas is mixed with the fourth stream of hydrogen in the fourth-stage mixer. It then enters the fourth-stage reactor at 152°C and 1.55 MPaG. The outlet gas is then cooled by the C6 component and mixed with the fifth stream of hydrogen in the fifth-stage mixer. It then enters the fifth-stage reactor at 152°C and 1.50 MPaG. The reaction products enter the C6 removal tower at 189°C and 1.5 MPaG. The distance between each stage mixer and its corresponding feed location for each reaction stage is 1 meter.

[0096] The deweight removal tower is a plate tower, and the operating temperature of the C6 removal tower is 180℃ and the operating pressure is 80kPaA; the temperature of the C6 component sampled from the side line of the C6 removal tower is 80℃.

[0097] The unit has a total hydrogen molar ratio of 1:2.5 to benzene in the C6 component, a total energy consumption of 175 kg standard oil / t hydrogenated terphenyl, a total fresh benzene conversion rate of 81%, and a total selectivity of 91%.

[0098]

Example 4

[0099] A 10,000-ton / year benzene hydrogenation alkylation to terphenyl hydrogenation unit adopts... Figure 1 The process technology involves a five-stage reactor, all of which are adiabatic fixed-bed reactors. The fresh benzene flow rate is 1800 kg / h, the hydrogen flow rate for each stage is 11.5 kg / h, and the C6 component flow rate from the C6 removal tower is 4800 kg / h.

[0100] The C6 component passes through interstage coolers, mixes with the first hydrogen stream in the first-stage mixer, and then enters the first-stage reactor at 180°C and 2 MPaG. The outlet gas temperature of the first-stage reactor is 219°C. After being cooled by the C6 component, the outlet gas is mixed with the second hydrogen stream in the second-stage mixer and then enters the second-stage reactor at 181°C and 1.95 MPaG. The outlet gas temperature of the second-stage reactor is 219°C. After being cooled by the C6 component, the outlet gas is mixed with three hydrogen streams in the third-stage mixer and then... The gas enters the third-stage reactor at 1℃ and 1.90 MPaG. The outlet gas temperature of the third-stage reactor is 220℃. After being cooled by the C6 component, the outlet gas is mixed with the fourth stream of hydrogen in the fourth-stage mixer. It then enters the fourth-stage reactor at 182℃ and 1.85 MPaG. After being cooled by the C6 component, the outlet gas is mixed with the fifth stream of hydrogen in the fifth-stage mixer. It then enters the fifth-stage reactor at 183℃ and 1.80 MPaG. The reaction products enter the C6 removal tower at 221℃ and 1.75 MPaG. The distance between each stage mixer and its corresponding feed location for each reaction stage is 1 meter.

[0101] The deweight removal tower is a plate tower, and the operating temperature of the C6 removal tower is 180℃ and the operating pressure is 80kPaA; the temperature of the C6 component sampled from the side line of the C6 removal tower is 80℃.

[0102] The unit has a total hydrogen to benzene molar ratio of 1:1.5, a total energy consumption of 170 kg standard oil / t hydrogenated terphenyl, a total fresh benzene conversion rate of 80%, and a total selectivity of 90%.

[0103] Comparative Example 1

[0104] The process of Example 2 was repeated, except that the fresh benzene was not passed through the C6 removal tower, but was mixed with the C6 component collected from the side stream of the C6 removal tower before entering the cooler E102. All other conditions remained the same.

[0105] The reactor was unstable and the temperature rise of the catalyst bed was uneven. After one month of operation, the catalyst was deactivated due to moisture brought by fresh benzene, and the unit was shut down to replace the catalyst.

[0106] Comparative Example 2

[0107] The process of Example 2 is repeated, except that the side stream exit point of the C6 stripping tower is at the same horizontal level (i.e., at the same height) as the fresh benzene feed point. All other conditions remain unchanged.

[0108] The C6 removal tower was unstable, the feed tray was frequently evacuated, and the extracted C6 component still contained 500-1000 ppm of moisture. The reactor was unstable, the catalyst bed temperature rose unevenly, and the catalyst was deactivated due to moisture brought by fresh benzene after three months of operation, so the unit was shut down to replace the catalyst.

[0109] The present invention has been described in detail above with reference to specific embodiments and exemplary examples; however, these descriptions should not be construed as limiting the present invention. Those skilled in the art will understand that various equivalent substitutions, modifications, or improvements can be made to the technical solutions and embodiments of the present invention without departing from the spirit and scope of the invention, and all such modifications and improvements fall within the scope of the present invention. The scope of protection of the present invention is defined by the appended claims.

Claims

1. A method of making a hydrogenated terphenyl comprising: The N-stage reaction is carried out with hydrogen and C6 components from a C6 removal column as raw materials, N≥2, and the reaction product of the Nth stage reaction enters the C6 removal column; each of the N-stage reactions is independently carried out in the presence of a catalyst, which is a metal-supported bifunctional solid acid hydrogenation alkylation catalyst; before the reaction product of the previous stage reaction enters the next stage reaction, the reaction product of the previous stage reaction is subjected to cooling treatment, wherein the C6 components from the C6 removal column are used as the cooling medium; in the C6 removal column, fresh benzene is used as the feed, the C6 components are taken out from the side line, overhead materials are taken out, and a product containing hydrogenated terphenyl is taken out from the bottom; in the C6 removal column, the feed position of the reaction product of the N-stage reaction is below the feed position of fresh benzene, and the feed position of fresh benzene is below the side line taking-out position of the C6 components.

2. The method of claim 1, wherein, The method comprises N-stage reactions, hydrogen is independently fed into each stage reaction in N streams, and the C6 components from the C6 removal column are fed into each stage reaction in one stream, and N = 2-6.

3. The method of claim 1, wherein, N is 2-4.

4. The method of claim 1, wherein, Before the C6 components enter each stage reaction, they are sequentially subjected to heat exchange treatment with the reaction products of the (N-1)th stage reaction, the (N-2)th stage reaction, …, and the 1st stage reaction, and then fed into each stage reaction in one stream after heat exchange.

5. The method of claim 1, wherein, the total amount of hydrogen used and the molar ratio of benzene in the C6 components are 1:(1.1-6); and / or, the hydrogen feed amount in each stage reaction is the same or different.

6. The method of claim 1, wherein, the total amount of hydrogen used and the molar ratio of benzene in the C6 components are 1:(1.5-4.5); and / or, the hydrogen feed amount in each stage reaction is the same.

7. The method of claim 1, wherein, Before entering each stage reaction, hydrogen is mixed with the C6 components or hydrogen is mixed with the reaction product of the previous stage reaction in a mixer.

8. The method of claim 7, wherein, The distance between each mixer and the feed position of the corresponding each stage reaction is greater than or equal to 0.5 meters.

9. The method of claim 7, wherein, The distance between each mixer and the feed position of the corresponding each stage reaction is greater than or equal to 1 meter.

10. The method of claim 1, wherein, The reaction inlet temperature of each stage reaction is independently 130-280°C; and the pressure of each stage reaction is independently 0.5-2.5 MPaG.

11. The method of claim 1, wherein, The reaction inlet temperature of each stage reaction is independently 150-220°C; and the pressure of each stage reaction is independently 1.0-2.0 MPaG.

12. The method of claim 1, wherein, the operating temperature of the C6 removal column is 150-250°C; the operating pressure is 20-200 kPaA; and / or, the temperature of the C6 components taken out from the side line of the C6 removal column is 60-120°C.

13. The method of claim 1, wherein, the operating temperature of the C6 removal column is 180-220°C; the operating pressure is 80-140 kPaA; and / or, the temperature of the C6 components taken out from the side line of the C6 removal column is 60-120°C.

14. The method according to one of claims 1 to 13, characterized in that The reaction product of the N-stage reaction enters from the lower part of the C6 removal column, and the fresh benzene enters from the upper part of the C6 removal column.

15. The method of claim 1, wherein, The catalyst comprises a carrier, a binder and an active component, wherein the carrier is a microporous organic silicon zeolite, the active component comprises a combination of a noble metal active component and a non-noble metal, and the binder is at least one selected from zinc oxide, aluminum oxide or titanium oxide.

16. The method of claim 15, wherein, the active component is at least one selected from palladium, ruthenium or platinum, and the non-noble metal is at least one selected from nickel, copper or cobalt.

17. The method of claim 15, wherein, The content of the carrier is 30-90 wt%, the content of the binder is 8-60 wt%, the content of the noble metal active component is 0.05-5 wt%, and the content of the non-noble metal active component is 10-60 wt%, based on 100 wt% of the total weight of the catalyst.

18. A system for producing hydrogenated terphenyls for carrying out the process of any one of claims 1 to 17, said system comprising N stages of reactors in series and a C6 removal column, wherein, The N-stage series reactors are N reactors connected in series, N = 2-6; and / or, The N-stage series reactors are N reactors connected in series, N = 2-6; and / or, A feed inlet is arranged at the upper part or top of each reactor, and a discharge outlet is arranged at the lower part or bottom of each reactor.

20. The system of claim 19, wherein, The N-stage series reactors are N reactors connected in series, N = 2-4; and / or, The discharge outlet of the upper-stage reactor is connected to the feed inlet of the lower-stage reactor.

21. The system of claim 18, wherein, The reaction material inlet is arranged at the lower part of the de-C6 column, and the fresh benzene inlet and the side-line outlet are arranged at the upper part of the de-C6 column, with the side-line outlet arranged above the fresh benzene inlet.

22. The system of claim 21, wherein, The reaction material inlet of the de-C6 column is connected to the discharge outlet of the N-stage reactor; and / or, the overhead outlet of the de-C6 column is arranged at the column top; and / or, the column bottom outlet of the de-C6 column is arranged at the column bottom. A mixer is independently arranged before the feed inlet of each reactor.

23. The system of any of claims 18-22, wherein, The distance between each mixer and the feed inlet of the corresponding reactor is greater than or equal to 0.5 meters.

24. The system of claim 23, wherein, The distance between each mixer and the feed inlet of the corresponding reactor is greater than or equal to 1 meter.

25. The system of claim 23, wherein, ​

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