Poly-alpha-olefin synthetic oil preparation system and method

Abstract

The invention provides a poly-alpha-olefin synthetic oil preparation system and a poly-alpha-olefin synthetic oil preparation method. The system comprises polymerization reaction equipment, stirring and mixing equipment, layering equipment, a monomer separation unit, a hydrofining unit and a product fractionation unit, a feeding hole of the stirring and mixing equipment is connected with a discharging hole of the polymerization reaction equipment, a discharging hole of the stirring and mixing equipment is connected with a feeding hole of the layering equipment, and an absorbent inlet of the stirring and mixing equipment is respectively connected with an absorbent source and a poly-alpha-olefin synthesis oil layer outlet of the layering equipment; a poly-alpha-olefin synthetic oil layer outlet of the layering equipment is connected with a feeding hole of the monomer separation unit, a poly-alpha-olefin synthetic oil outlet of the monomer separation unit is connected with a feeding hole of the hydrofining unit, and a poly-alpha-olefin synthetic oil outlet of the hydrofining unit is connected with a feeding hole of the product fractionation unit; wherein a fluid discharge pipeline is arranged on a connecting pipeline between an absorbent inlet of the stirring and mixing equipment and a poly-alpha-olefin synthesis oil layer outlet of the layering equipment.

Description

Technical Field

[0001] The present invention relates to a system and method for preparing poly-α-olefin synthetic oil. Background Art

[0002] Class IV lubricant base oil - poly-α-olefin synthetic oil is currently the synthetic lubricant base oil with the best performance, having excellent properties such as high viscosity index, low volatility, and good low-temperature fluidity, and can be applicable to various harsh working conditions. It is not only widely used in the civilian industry but also the main source of special lubricating base oil for industries such as aviation, aerospace, and military. The global demand for poly-α-olefin synthetic oil has always shown a trend of supply falling short of demand. However, in the process of producing poly-α-olefin synthetic oil, a catalyst system containing BF 3 is usually required. BF 3 in poly-α-olefin synthetic oil will cause corrosion of equipment and pipelines, bringing many inconveniences to enterprises, and will also reduce the quality of poly-α-olefin synthetic oil. Therefore, it is necessary to remove BF 3 from poly-α-olefin synthetic oil. Currently, the commonly used method for removing BF 3 in poly-α-olefin synthetic oil is caustic washing. However, caustic washing will generate a large amount of fluorine-containing and oil-containing wastewater, increasing the burden on the sewage treatment plant.

[0003] Therefore, some researchers have tried to use other methods to remove BF 3 from poly-α-olefin synthetic oil. However, these methods all have many problems. For example, the method of reducing the pressure of the polymerization product to flash off the excessive boron trifluoride for BF 3 removal, and for another example, the method of gas-liquid separation after heating for BF 3 removal. These methods have a low BF 3 removal rate and usually still need to be combined with caustic washing or alkali addition and neutralization treatment, and are prone to cause BF 3 to accumulate in the separation equipment and corrode the equipment.

[0004] In summary, there is an urgent need to study a poly-α-olefin synthetic oil preparation technology that can avoid the generation of fluorine-containing wastewater and reduce the corrosion of pipelines by BF 3 Summary of the Invention

[0005] The purpose of the present invention is to provide a technical solution for preparing poly-α-olefin synthetic oil that can avoid the generation of fluorine-containing wastewater and reduce the corrosion of pipelines by BF 3

[0006] To solve the above problems, the present invention provides the following two aspects of technical solutions.

[0007] ​​In a first aspect, the present invention provides a polyalphaolefin synthetic oil preparation system, wherein the system includes a polymerization reaction device, a stirring and mixing device, a stratification device, a monomer separation unit, a hydrorefining unit, and a product fractionation unit;

[0008] The feed inlet of the stirring and mixing device is connected to the discharge outlet of the polymerization reaction device, the discharge outlet of the stirring and mixing device is connected to the feed inlet of the stratification device, the absorbent inlets of the stirring and mixing device are respectively connected to the adsorbent source and the polyalphaolefin synthetic oil layer outlet of the stratification device, the polyalphaolefin synthetic oil layer outlet of the stratification device is connected to the feed inlet of the monomer separation unit, the polyalphaolefin synthetic oil outlet of the monomer separation unit is connected to the feed inlet of the hydrorefining unit, and the polyalphaolefin synthetic oil outlet of the hydrorefining unit is connected to the feed inlet of the product fractionation unit; wherein, a fluid discharge pipeline is provided on the pipeline connecting the absorbent inlet of the stirring and mixing device and the polyalphaolefin synthetic oil layer outlet of the stratification device.

[0009] In the above polyalphaolefin synthetic oil preparation system, a stirring and mixing device and a stratification device are provided between the polymerization reaction device and the monomer separation unit. With the cooperation of the stirring and mixing device and the stratification device, it is possible to effectively and continuously remove the catalyst system containing BF 3 in the synthetic oil during the synthesis of polyalphaolefin, and no fluorine-containing wastewater will be generated. Then, the monomer separation unit, the hydrorefining unit, and the product fractionation unit are sequentially used to purify, refine, and fractionate the polyalphaolefin synthetic oil from which the catalyst system containing BF 3 has been removed, to obtain high-quality PAO base oil products with different viscosities.

[0010] According to a preferred embodiment of the first aspect, the polymerization reaction device is a polymerization reactor.

[0011] According to a preferred embodiment of the first aspect, the polymerization reaction device is provided with a feed inlet, a BF 3 inlet, a solid-liquid catalyst inlet, and a discharge outlet; wherein, the feed inlet and the solid-liquid catalyst inlet are provided at the bottom of the polymerization reaction device, and the BF 3 inlet is provided at the top of the polymerization reaction device;

[0012] More preferably, a stirrer is provided inside the polymerization reaction device;

[0013] More preferably, the polymerization reaction device is provided with a temperature control part, and the temperature control part is provided with a heat transfer medium inlet and a heat transfer medium outlet.

[0014] According to a preferred embodiment of the first aspect, the polyalphaolefin synthetic oil preparation system further includes a drying device, and the discharge outlet of the drying device is connected to the feed inlet of the polymerization reaction device.

[0015] According to a preferred embodiment of the first aspect, the monomer outlet of the monomer separation unit is connected to the feed inlet of the polymerization reaction equipment.

[0016] According to a preferred embodiment of the first aspect, a sampling point is provided on the connecting pipeline between the absorbent inlet of the stirring and mixing equipment and the poly-α-olefin synthetic oil layer outlet of the stratifying equipment.

[0017] According to a preferred embodiment of the first aspect, the monomer separation unit includes a heating device and a monomer stripping tower; the feed inlet of the heating device is connected to the poly-α-olefin synthetic oil layer outlet of the stratifying equipment, and the discharge outlet of the heating device is connected to the feed inlet of the monomer stripping tower;

[0018] More preferably, the monomer separation unit further includes a gas-liquid separation device; the gas outlet of the monomer stripping tower is connected to the feed inlet of the gas-liquid separation device, and the liquid outlet of the gas-liquid separation device serves as the monomer outlet of the monomer separation unit.

[0019] According to a preferred embodiment of the first aspect, the hydrofining unit includes a hydrogenation reaction device, a hydrogen separation device, and an impurity stripping tower; the feed inlet of the hydrogenation reaction device is connected to the poly-α-olefin synthetic oil outlet of the monomer separation unit, the discharge outlet of the hydrogenation reaction device is connected to the feed inlet of the hydrogen separation device, the liquid phase outlet of the hydrogen separation device is connected to the feed inlet of the impurity stripping tower, and the poly-α-olefin synthetic oil outlet of the impurity stripping tower serves as the poly-α-olefin synthetic oil outlet of the hydrofining unit;

[0020] More preferably, the hydrogen separation device includes a high-pressure separator and a low-pressure separator; the feed inlet of the high-pressure separator is connected to the discharge outlet of the hydrogenation reaction device, the liquid phase outlet of the high-pressure separator is connected to the feed inlet of the low-pressure separator, and the liquid phase outlet of the low-pressure separator is connected to the feed inlet of the impurity stripping tower; further, the gas phase outlet of the high-pressure separator is connected to the hydrogenation reaction device to realize the recycling of the hydrogen separated by the high-pressure separator back to the hydrogenation reaction device for use.

[0021] More preferably, the impurity stripping tower is a plate column or a packed column;

[0022] More preferably, the number of theoretical plates of the impurity stripping tower is 3 - 25, the poly-α-olefin synthetic oil outlet of the impurity stripping tower is arranged at the bottom of the impurity stripping tower, and the gas stripping outlet of the impurity stripping tower is arranged at the top of the impurity stripping tower;

[0023] More preferably, the hydrogenation reaction device is selected from one or 2 - 4 series-connected fixed-bed hydrogenation reactors.

[0024] According to a preferred embodiment of the first aspect, the product fractionation unit includes an atmospheric column and a vacuum column. The feed inlet of the atmospheric column is connected to the outlet of the synthetic polyalphaolefin oil of the hydrofining unit. The bottom material outlet of the atmospheric column is connected to the feed inlet of the vacuum column. The side line and bottom material outlets of the vacuum column serve as the outlets for PAO base oil products with different viscosities.

[0025] More preferably, the atmospheric column is a plate column or a packed column.

[0026] More preferably, the vacuum column is a plate column or a packed column.

[0027] More preferably, the number of theoretical plates of the atmospheric column is 20 - 40, and the feed plate position is 15 - 30.

[0028] More preferably, the number of theoretical plates of the vacuum column is 35 - 65, and the feed plate position is 25 - 50.

[0029] In a second aspect, the present invention provides a method for preparing synthetic polyalphaolefin oil. This method is carried out using the synthetic polyalphaolefin oil preparation system provided in the first aspect. This method includes:

[0030] Feeding an α-olefin raw material oil and a catalyst system containing BF 3 into a polymerization reaction device, and carrying out the polymerization reaction of the α-olefin raw material oil under the catalyst system containing BF 3 in the polymerization reaction device; the catalyst system containing BF 3 includes gaseous BF 3 and a catalyst.

[0031] Feeding the polymerization reaction product in the polymerization reaction device, an alcohol absorbent, and the absorbent layer conveyed out from a separation device into a stirring and mixing device, and carrying out mixing in the stirring and mixing device; wherein, BF 3 complexes with the alcohol absorbent to form a complex product of BF 3 and the alcohol.

[0032] Feeding the product after mixing in the stirring and mixing device into a separation device for separation to obtain an absorbent layer and a synthetic polyalphaolefin oil layer, and conveying the absorbent layer and the synthetic polyalphaolefin oil layer out from the separation device respectively; wherein, the catalyst system containing BF 3 exists in the absorbent layer, and the complex product of BF 3 and the alcohol and the catalyst are dissolved in the absorbent layer; wherein, a part of the absorbent layer conveyed out from the separation device is conveyed to the stirring and mixing device.

[0033] Conveying the synthetic polyalphaolefin oil layer conveyed out from the separation device to a monomer separation unit to carry out α-olefin monomer separation to obtain a crude synthetic polyalphaolefin oil product.

[0034] The crude poly-α-olefin synthetic oil product in the monomer separation unit is transported to the hydrofining unit for hydrofining;

[0035] The hydrofined product in the hydrofining unit is transported to the product fractionation unit for fractionation to obtain the PAO base oil product.

[0036] According to a preferred embodiment of the second aspect, wherein the α-olefin feedstock oil and the catalyst system containing BF 3 are transported to the polymerization reaction equipment, and the polymerization reaction of the α-olefin feedstock oil under the catalyst system containing BF 3 is achieved by the following method:

[0037] The α-olefin feedstock oil and the catalyst are transported to the polymerization reaction equipment from the bottom of the polymerization reaction equipment, and at the same time, the gaseous BF 3 is transported to the polymerization reaction equipment from the top of the polymerization reaction equipment for pressure regulation, and the polymerization reaction of the α-olefin feedstock oil under the catalyst system containing BF 3 is carried out in the polymerization reaction equipment.

[0038] According to a preferred embodiment of the second aspect, wherein the catalyst includes at least one of a monohydric alcohol and a dihydric alcohol; more preferably, the catalyst includes at least one of a monohydric alcohol with a carbon number not exceeding 10 and a dihydric alcohol with a carbon number not exceeding 10.

[0039] According to a preferred embodiment of the second aspect, wherein the temperature of the polymerization reaction is 10 - 100 °C, the pressure is 0.05 - 2.0 MPaG, and the time is 0.2 - 40 h.

[0040] According to a preferred embodiment of the second aspect, wherein the mass ratio of the α-olefin feedstock oil, BF 3 , and the catalyst is: 100:0.0001 - 0.5:0.1 - 1.

[0041] According to a preferred embodiment of the second aspect, wherein the method further includes: detecting the alcohol absorbent content of the absorbent layer transported out of the stratification equipment to ensure that all BF 3 is used for complexation with the alcohol absorbent.

[0042] According to a preferred embodiment of the second aspect, wherein the mass ratio of the alcohol absorbent transported to the stirring and mixing equipment to the BF 3 transported to the polymerization reaction equipment is: 4 - 40:1.

[0043] According to a preferred embodiment of the second aspect, wherein the alcohol absorbent includes at least one of monohydric alcohols with a carbon number of 1 - 10.

[0044] According to a preferred embodiment of the second aspect, at least one of α-olefins having 4 to 14 carbon atoms is included in the α-olefin feedstock oil;

[0045] More preferably, in the α-olefin feedstock oil, the mass content of α-olefins having 4 to 14 carbon atoms is ≥80%.

[0046] According to a preferred embodiment of the second aspect, the mass content of normal α-olefins in the α-olefin feedstock oil is ≥60%.

[0047] According to a preferred embodiment of the second aspect, the method further includes: drying the α-olefin feedstock oil before it is transported to the polymerization reaction equipment.

[0048] According to a preferred embodiment of the second aspect, transporting the poly-α-olefin synthetic oil layer transported out of the stratification equipment to the monomer separation unit for α-olefin monomer separation includes:

[0049] heating the poly-α-olefin synthetic oil layer transported out of the stratification equipment to 100 - 300 °C and then transporting it to the monomer stripping tower, and carrying out α-olefin monomer separation in the monomer stripping tower by means of nitrogen stripping;

[0050] More preferably, the separated gas enters the gas-liquid separation equipment to be cooled and then undergoes gas-liquid separation, and the separated liquid phase is the α-olefin monomer; even more preferably, the α-olefin monomer is transported to the polymerization reaction equipment for recycling;

[0051] More preferably, the temperature of the monomer stripping tower is 10 - 200 °C, the pressure is 5 KPa - 150 KPa (absolute pressure), the mass ratio of gas to oil is 0.01 - 0.5, and the injection pressure of nitrogen is 0.10 - 2 MPa (gauge pressure).

[0052] According to a preferred embodiment of the second aspect, transporting the crude poly-α-olefin synthetic oil in the monomer separation unit to the hydrofining unit for hydrofining includes:

[0053] boosting the pressure of the crude poly-α-olefin synthetic oil output from the monomer separation unit to 2 - 6 MPa (gauge pressure), mixing it with hydrogen, heating it to 150 - 300 °C, and carrying out a hydrogenation reaction in a hydrogenation reaction equipment; transporting the hydrogenation reaction product in the hydrogenation reaction equipment to a hydrogen separator for hydrogen separation; transporting the liquid phase product after hydrogen separation in the hydrogen separator to an impurity stripping tower for nitrogen stripping, and the obtained liquid phase product is the hydrofined product;

[0054] More preferably, the overall volume space velocity of the hydrogenation reaction is 0.1 - 1 h -1 , and the standard volume ratio of hydrogen to oil is 10:1 - 1000:1;

[0055] More preferably, during the nitrogen stripping process, the temperature of the nitrogen used is 10 - 200 °C, the pressure is 0.10 - 2 MPa (gauge pressure), and the mass ratio of gas to oil is 50 - 100:1;

[0056] More preferably, the pressure of the impurity stripping column is 0 - 1 MPa (gauge pressure), the top temperature is 10 - 200 °C, and the bottom temperature is 10 - 200 °C;

[0057] More preferably, the transportation of the hydrogenation reaction product in the hydrogenation reaction equipment to the hydrogen separator for hydrogen separation includes:

[0058] Transporting the hydrogenation reaction product in the hydrogenation reaction equipment to the high-pressure separator for hydrogen separation, and the liquid-phase product after hydrogen separation in the high-pressure separator is transported to the low-pressure separator for hydrogen separation, and the liquid-phase product after hydrogen separation in the low-pressure separator is used as the liquid-phase product after hydrogen separation in the hydrogen separator;

[0059] More preferably, the gas-phase product after hydrogen separation in the high-pressure separator is mixed with the product after the crude poly-α-olefin synthetic oil output from the monomer separation unit is boosted to 2 - 6 MPa (gauge pressure).

[0060] According to the preferred embodiment of the second aspect, wherein, transporting the hydrorefined product in the hydrorefining unit to the product fractionation unit for fractionation to obtain the PAO base oil product includes:

[0061] Heating the hydrorefined product output from the hydrorefining unit to 300 - 380 °C and then transporting it to the atmospheric column for atmospheric fractionation; heating the bottom material output from the atmospheric column to 300 - 380 °C and then transporting it to the vacuum column for vacuum fractionation to obtain the PAO base oil product;

[0062] More preferably, the top pressure of the atmospheric column is 0.01 - 0.1 MPa (gauge pressure), the top temperature is 200 - 300 °C, the bottom temperature is 250 - 380 °C, the reflux ratio is 0.1 - 20, and the mass ratio of steam injection at the bottom of the column to the feed is 0.01 - 1;

[0063] More preferably, the top pressure of the vacuum column is 5 - 50 KPa (absolute pressure), the top temperature is 40 - 100 °C, the bottom temperature is 100 - 380 °C, and the mass ratio of steam injection at the bottom of the column to the feed is 0.01 - 1.

[0064] Compared with the prior art, the technical solution provided by the present invention has the following beneficial effects:

[0065] 1. The technical solution provided by the present invention can realize the removal of catalysts and BF in the α-olefin polymerization product 3Carry out "harmless" separation to solve the difficult problems of wastewater and waste residue treatment in traditional water caustic washing process and solid adsorption process.

[0066] 2. The preferred mode of the technical solution provided by the present invention can further achieve efficient recycling on the basis of realizing the "harmless" separation of the catalyst and BF in the α-olefin polymerization product, thereby improving the technical economy. 3 On the basis of carrying out "harmless" separation, it can be further recycled efficiently to improve the technical economy.

[0067] 3. The technical solution provided by the present invention uses an alcohol absorbent to complex and absorb the gaseous BF dissolved in the α-olefin polymerization product. 3 to achieve zero emission of BF 3 gas, which is pollution-free to the environment and can reduce the corrosion of the device.

[0068] 4. The technical solution provided by the present invention integrates the whole process technologies of continuous polymerization, catalyst separation, deep hydrogenation of base oil and fractionation of base oil products with different viscosities, with simple process, complete process flow, low operation difficulty, high production efficiency, advanced technology and easy promotion. BRIEF DESCRIPTION OF THE DRAWINGS

[0069] Figure 1 It is a schematic structural diagram of the poly-α-olefin synthetic oil preparation system in Example 1. DETAILED DESCRIPTION OF THE EMBODIMENTS

[0070] In order to have a clearer understanding of the technical features, objectives and beneficial effects of the present invention, the technical solution of the present invention will be described in detail below, but it should not be construed as a limitation on the scope of implementation of the present invention.

[0071] Example 1

[0072] This example provides a poly-α-olefin synthetic oil preparation system, as Figure 1 shown, the system includes a drying device 1, a polymerization reaction device 2, a stirring and mixing device 3, a stratifying device 4, a monomer separation unit, a hydrorefining unit and a product fractionation unit.

[0073] The polymerization reaction device 2 is provided with a feed inlet, a BF 3 inlet, a solid-liquid phase catalyst inlet and a discharge outlet; the feed inlet and the solid-liquid phase catalyst inlet are arranged at the bottom of the polymerization reaction device 2, and the BF 3 inlet is arranged at the top of the polymerization reaction device 2; a stirrer is arranged in the polymerization reaction device 2; the polymerization reaction device 2 is provided with a temperature control part, and the temperature control part is provided with a heat transfer medium inlet and a heat transfer medium outlet.

[0074] The monomer separation unit includes a heating device 5, a monomer stripping tower 6 and a gas-liquid separation device 7.

[0075] The hydrofining unit includes a hydroreaction device 8, a high-pressure separator 9, a low-pressure separator 10, an impurity stripping column 11, and a recycle hydrogen compressor 12.

[0076] The product fractionation unit includes an atmospheric column 13, a vacuum column 14, an oil-water separation tank 15, a heating furnace 16, an oil-water separation tank 17, and a heating furnace 18.

[0077] The discharge port of the drying device 1 is connected to the feed port of the polymerization reaction device 2. The feed port of the stirring and mixing device 3 is connected to the discharge port of the polymerization reaction device 2. The discharge port of the stirring and mixing device 3 is connected to the feed port of the stratification device 4. The absorbent inlets of the stirring and mixing device 3 are respectively connected to the adsorbent source and the poly-α-olefin synthetic oil layer outlet of the stratification device 4. The poly-α-olefin synthetic oil layer outlet of the stratification device 4 is connected to the feed port of the heating device 5. The discharge port of the heating device 5 is connected to the feed port of the monomer stripping column 6. The gas outlet of the monomer stripping column 6 is connected to the feed port of the gas-liquid separation device 7. The liquid outlet of the gas-liquid separation device 7 is connected to the feed port of the polymerization reaction device 2. Among them, a fluid discharge pipeline and a sampling point are provided on the connecting pipeline between the absorbent inlet of the stirring and mixing device 3 and the poly-α-olefin synthetic oil layer outlet of the stratification device 4. The poly-α-olefin synthetic oil outlet of the monomer stripping column 6 is connected to the feed port of the hydroreaction device 8. The discharge port of the hydroreaction device 8 is connected to the feed port of the high-pressure separator 9. The liquid phase outlet of the high-pressure separator 9 is connected to the feed port of the low-pressure separator 10. The liquid phase outlet of the low-pressure separator 10 is connected to the feed port of the impurity stripping column 11. The gas phase outlet of the high-pressure separator 9 is connected to the feed port of the recycle hydrogen compressor 12. The material outlet of the recycle hydrogen compressor 12 is connected to the connecting pipeline between the poly-α-olefin synthetic oil outlet of the monomer stripping column 6 and the feed port of the hydroreaction device 8. The poly-α-olefin synthetic oil outlet of the impurity stripping column 11 is connected to the feed port of the heating furnace 16. The material outlet of the heating furnace 16 is connected to the feed port of the atmospheric column 13. The bottom material outlet of the atmospheric column 13 is connected to the feed port of the heating furnace 18. The material outlet of the heating furnace 18 is connected to the feed port of the vacuum column 14. The top material outlet of the atmospheric column 13 is connected to the feed port of the oil-water separation tank 15. The oil phase outlet of the oil-water separation tank 15 is respectively connected to the discharge pipeline and the atmospheric column 13. The top material outlet of the vacuum column 14 is connected to the feed port of the oil-water separation tank 17. The oil phase outlet of the oil-water separation tank 17 is connected to the discharge pipeline. The side lines and the bottom material outlet of the vacuum column are used as the outlets of PAO base oil products with different viscosities.

[0078] Among them, the polymerization reaction device 2 is a polymerization reactor.

[0079] Among them, the drying device 2 is a parallel-connected molecular sieve drying tower.

[0080] Among them, the gas-liquid separation device 7 is a gas-liquid separation tank.

[0081] Among them, the stirring and mixing device 3 is a stirring tank.

[0082] Among them, the monomer stripping tower 6 is a plate tower or a packed tower.

[0083] Among them, the impurity stripping tower is a plate tower or a packed tower;

[0084] Among them, the number of theoretical plates of the impurity stripping tower is 3 - 25. The poly-α-olefin synthetic oil outlet of the impurity stripping tower is arranged at the bottom of the impurity stripping tower, and the stripping gas outlet of the impurity stripping tower is arranged at the top of the impurity stripping tower.

[0085] Among them, the atmospheric tower is a plate tower or a packed tower;

[0086] Among them, the vacuum tower is a plate tower or a packed tower;

[0087] Among them, the number of theoretical plates of the atmospheric tower is 20 - 40, and the feed plate position is 15 - 30;

[0088] Among them, the number of theoretical plates of the vacuum tower is 35 - 65, and the feed plate position is 15 - 30.

[0089] Example 2

[0090] This example provides a method for preparing poly-α-olefin synthetic oil. This method is carried out using the poly-α-olefin synthetic oil preparation system provided in Example 1. This method includes:

[0091] Step 1: The α-olefin raw oil stream A at a temperature of 25°C and a pressure of 0.05 MPa enters the drying device 1 at a flow rate of 500 kg / h for drying and water removal.

[0092] Step 2: The α-olefin raw oil stream B (with a water content of 10 μg / g) dried by the drying device 1 is mixed with the liquid-phase α-olefin monomer stream J separated by the gas-liquid separation device 7 and is fed into the polymerization reaction device 2 from the bottom of the polymerization reaction device 2. At the same time, the catalyst stream C (methanol) enters the polymerization reaction device 2 from the bottom of the polymerization reaction device 2 at a flow rate of 2.4 kg / h. At the same time, the gas-phase BF 3 stream D enters the polymerization reaction device 2 from the top of the polymerization reaction device 2 at a flow rate of 0.78 kg / h, and an α-olefin polymerization reaction is carried out in the polymerization reaction device 2; among them, the temperature of the polymerization reaction device 2 is 10°C, the pressure is 0.05 MPaG, and the residence time of the material in the polymerization reaction device 2 is 5 h.

[0093] Step 3: The polymerization reaction product stream E (with a flow rate of approximately 599.9 kg / h during the normal operation stage) in the polymerization reaction apparatus 2 enters the stirring and mixing apparatus 3. Meanwhile, the alcohol absorbent stream F (methanol is used in this embodiment) at a temperature of 25°C and a pressure of 0.4 MPa (gauge pressure) enters the stirring apparatus 3 at a flow rate of 2.6 kg / h. At the same time, the absorbent layer stream G at a temperature of 30°C and a pressure of 0.8 MPa (gauge pressure) (with a flow rate of approximately 3.2 kg / h during the normal operation stage) conveyed from the separation apparatus 4 is conveyed to the stirring and mixing apparatus 3 and is mixed evenly therein. Among them, BF 3 forms a complex with the alcohol absorbent to form BF 3 a complex product with the alcohol;

[0094] The product evenly mixed in the stirring and mixing apparatus 3 is conveyed to the separation apparatus 4 for separation to obtain an absorbent layer and a poly-α-olefin synthetic oil layer, and the absorbent layer and the poly-α-olefin synthetic oil layer are respectively conveyed out of the separation apparatus; among them, the catalyst system containing BF 3 exists in the absorbent layer, and BF 3 the BF formed by complexing with the alcohol absorbent 3 the complex product with the alcohol and the catalyst are dissolved in the alcohol absorbent; among them, a part of the absorbent layer conveyed out of the separation apparatus 4 is conveyed to the stirring and mixing apparatus 3 as the absorbent layer stream G.

[0095] Step 4: The poly-α-olefin synthetic oil layer stream H conveyed out of the separation apparatus 4 is heated to 100°C by the heating apparatus 5 and then enters the middle section of the monomer stripping tower 6. Meanwhile, the nitrogen stream I enters the monomer stripping tower 6 from the bottom of the monomer stripping tower 6. In the monomer stripping tower 6, α-olefin monomer nitrogen stripping separation is carried out. The α-olefin monomer and nitrogen are taken out from the top of the monomer stripping tower 6 as a gas phase and enter the gas-liquid separation apparatus 7 for cooling and gas-liquid separation to obtain the liquid-phase α-olefin monomer stream J. The liquid-phase α-olefin monomer stream J is mixed with the α-olefin raw oil stream B (with a water content of 10 μg / g) dried by the drying apparatus 1 for recycling; the poly-α-olefin synthetic oil crude product stream K is output from the bottom of the monomer stripping tower 6.

[0096] Step 5: The crude poly-α-olefin synthetic oil product stream K output from the bottom of the monomer stripping tower 6 is pressurized to 2.0 MPa (gauge pressure) by a pump, mixed with the hydrogen stream L with a flow rate of 39 kg / h and the hydrogen output from the gas-phase outlet of the high-pressure separator 9, heated to 215 °C by a heating furnace, and then enters the hydrogenation reaction equipment 8 for hydrogenation reaction. The hydrogenation reaction product in the hydrogenation reaction equipment 8 is transported to the high-pressure separator 9 for hydrogen separation. The liquid-phase product after hydrogen separation in the high-pressure separator 9 is transported to the low-pressure separator 10 for hydrogen separation. The liquid-phase product after hydrogen separation in the low-pressure separator 10 is transported to the impurity stripping tower 11 for nitrogen stripping, and the obtained liquid-phase product is the hydrorefined product. Among them, the hydrogen output from the gas-phase outlet after hydrogen separation in the high-pressure separator 9 is pressurized to 5 MPa (gauge pressure) by the recycle hydrogen compressor 12 and then mixed with the hydrogen stream L with a flow rate of 39 kg / h and the product of the crude poly-α-olefin synthetic oil product stream K output from the bottom of the monomer stripping tower 6 pressurized to 2.0 MPa (gauge pressure); the gas-phase output from the gas-phase outlet after hydrogen separation in the low-pressure separator 10 is cooled and then discharged to the flare.

[0097] Among them, the temperature of the high-pressure separator 9 is 255 °C and the pressure is 4 MPa (gauge pressure);

[0098] Among them, the overall volume space velocity of the hydrogenation reaction is 0.1 h -1 ;

[0099] Among them, during the nitrogen stripping process, the temperature of the nitrogen used is 25 °C, the pressure is 0.4 MPa (gauge pressure), and the mass ratio of gas to oil is 100:1;

[0100] Among them, the pressure of the impurity stripping tower is 0.05 MPa (gauge pressure), the top temperature is 155 °C, the bottom temperature is 160 °C, the number of theoretical plates is 10, and the feed plate position is the top feed of the feedstock oil.

[0101] Step 6: The hydrorefined product output from the impurity stripping tower 11 is heated to 290 °C by the heating furnace 16 and then transported to the atmospheric tower 13 for atmospheric fractionation; the number of theoretical plates of the atmospheric tower 13 is 35, the feed plate position is 24, the top temperature is 100 °C, the bottom temperature is 270 °C, the reflux ratio is 2, and the mass ratio of steam injection at the bottom to the feed is 0.01. The PAO base oil product is obtained at the top of the atmospheric tower 13.

[0102] The bottom material output from the atmospheric column 13 is heated to 300 °C by the heating furnace 18 and then transported to the vacuum column 14 for vacuum fractionation. From the bottom to the top of the column, PAO base oil product streams M, N, O, and P are obtained; the theoretical number of plates of the vacuum column 14 is 60, the feed plate position is 45, the top pressure is 5 KPa (absolute pressure), the top temperature is 60 °C, the bottom temperature is 250 °C, the reflux ratio is 2, and the mass ratio of steam injection at the bottom to the feed is 0.01. The parameters and compositions of each logistics are shown in Table 1.

[0103] Table 1

[0104]

[0105]

[0106] Example 3

[0107] This example provides a method for preparing polyalphaolefin synthetic oil. This method is carried out using the polyalphaolefin synthetic oil preparation system provided in Example 1. This method includes:

[0108] Step 1: The α-olefin raw oil stream A at a temperature of 25 °C and a pressure of 0.05 MPa enters the drying equipment 1 at a flow rate of 1255.2 kg / h for drying and water removal.

[0109] Step 2: The α-olefin raw oil stream B (with a water content of 10 μg / g) dried by the drying equipment 1 is mixed with the liquid-phase α-olefin monomer stream J separated by the gas-liquid separation equipment 7 and fed into the polymerization reaction equipment 2 from the bottom of the polymerization reaction equipment 2. At the same time, the catalyst stream C (octanol) enters the polymerization reaction equipment 2 from the bottom of the polymerization reaction equipment 2 at a flow rate of 5.4 kg / h. At the same time, the gas-phase BF 3 stream D enters the polymerization reaction equipment 2 from the top of the polymerization reaction equipment 2 at a flow rate of 2.1 kg / h, and an α-olefin polymerization reaction is carried out in the polymerization reaction equipment 2; among them, the temperature of the polymerization reaction equipment 2 is 50 °C, the pressure is 1.0 MPaG, and the residence time of the material in the polymerization reaction equipment 2 is 10 h.

[0110] Step 3: The polymerization reaction product stream E (with a flow rate of about 1404.3 kg / h during the normal operation stage) in the polymerization reaction equipment 2 enters the stirring and mixing equipment 3. At the same time, the alcohol absorbent stream F (octanol is used in this example) at a temperature of 25 °C and a pressure of 0.4 MPa (gauge pressure) enters the stirring equipment 3 at a flow rate of 6.7 kg / h. At the same time, the absorbent layer stream G at a temperature of 40 °C and a pressure of 0.8 MPa (gauge pressure) (with a flow rate of about 12.2 kg / h during the normal operation stage) sent out from the stratification equipment 4 is transported to the stirring and mixing equipment 3 and mixed in the stirring and mixing equipment 3; among them, BF 3Complex with an alcohol absorbent to form BF 3 Complex product with alcohol;

[0111] Transfer the product mixed evenly in the stirring and mixing device 3 to the stratifying device 4 for stratification to obtain an absorbent layer and a poly-α-olefin synthetic oil layer, and convey the absorbent layer and the poly-α-olefin synthetic oil layer out of the stratifying device respectively; among them, the catalyst system containing BF 3 exists in the absorbent layer, and the BF 3 Complex of BF formed by complexing with an alcohol absorbent 3 Complex product with alcohol and the catalyst are dissolved in the alcohol absorbent; among them, a part of the absorbent layer conveyed out of the stratifying device 4 is conveyed to the stirring and mixing device 3 as the absorbent layer flow G.

[0112] Step 4: Heat the poly-α-olefin synthetic oil layer flow H conveyed out of the stratifying device 4 to 180 °C by the heating device 5 and then enter the middle section of the monomer stripping tower 6 from the middle section of the monomer stripping tower 6. At the same time, the nitrogen gas flow I enters the monomer stripping tower 6 from the bottom of the monomer stripping tower 6. In the monomer stripping tower 6, α-olefin monomer nitrogen stripping separation is carried out. The α-olefin monomer and nitrogen are taken out from the top of the monomer stripping tower 6 as a gas phase and enter the gas-liquid separation device 7 for cooling and gas-liquid separation to obtain the liquid-phase α-olefin monomer flow J. Mix the liquid-phase α-olefin monomer flow J with the α-olefin raw oil flow B (water content is 10 μg / g) dried by the drying device 1 for recycling; the poly-α-olefin synthetic oil crude product flow K is output from the bottom of the monomer stripping tower 6.

[0113] Step 5: The poly-α-olefin synthetic oil crude product flow K output from the bottom of the monomer stripping tower 6 is boosted to 4.5 MPa (gauge pressure) by a pump, mixed with the hydrogen gas flow L with a flow rate of 78 kg / h and the hydrogen gas output from the gas phase outlet of the high-pressure separator 9, heated to 215 °C by a heating furnace, and enter the hydrogenation reaction device 8 for hydrogenation reaction. The hydrogenation reaction product in the hydrogenation reaction device 8 is conveyed to the high-pressure separator 9 for hydrogen separation. The liquid-phase product after hydrogen separation in the high-pressure separator 9 is conveyed to the low-pressure separator 10 for hydrogen separation. The liquid-phase product after hydrogen separation in the low-pressure separator 10 is conveyed to the impurity stripping tower 11 for nitrogen stripping to obtain the liquid-phase product as the hydrorefined product. Among them, the hydrogen gas output from the gas phase outlet after hydrogen separation in the high-pressure separator 9 is boosted to 5 MPa (gauge pressure) by the recycle hydrogen compressor 12 and then mixed with the hydrogen gas flow L with a flow rate of 78 kg / h and the product after the poly-α-olefin synthetic oil crude product flow K output from the bottom of the monomer stripping tower 6 is boosted to 2.0 MPa (gauge pressure); the gas phase output from the gas phase outlet after hydrogen separation in the low-pressure separator 10 is discharged to the flare after cooling.

[0114] Among them, the temperature of the high-pressure separator 9 is 255 °C and the pressure is 4 MPa (gauge pressure);

[0115] Among them, the total volume space velocity of the hydrogenation reaction is 0.2 h -1 ;

[0116] Among them, during the nitrogen stripping process, the temperature of the nitrogen used is 25 °C, the pressure is 0.4 MPa (gauge pressure), and the mass ratio of gas to oil is 50:1;

[0117] Among them, the tower pressure of the impurity stripper is 0.1 MPa (gauge pressure), the top temperature of the tower is 180 °C, the bottom temperature of the tower is 210 °C, the number of theoretical plates is 16, and the feed plate position is the top feed of the feedstock oil tower.

[0118] Step 6: Heat the hydrofined product output from the impurity stripper 11 to 335 °C by the heating furnace 16 and then transport it to the atmospheric tower 13 for atmospheric fractionation; the number of theoretical plates of the atmospheric tower 13 is 25, the feed plate position is 18, the top temperature of the tower is 100 °C, the bottom temperature of the tower is 270 °C, the reflux ratio is 3, and the mass ratio of steam injection at the bottom of the tower to the feed is 0.5. The PAO base oil product is obtained at the top of the atmospheric tower 13.

[0119] Heat the bottom material output from the atmospheric tower 13 to 345 °C by the heating furnace 18 and then transport it to the vacuum tower 14 for vacuum fractionation. From the bottom to the top of the tower, PAO base oil product stream M, PAO base oil product stream N, PAO base oil product stream O, and PAO base oil product stream P are obtained; the number of theoretical plates of the vacuum tower 14 is 40, the feed plate position is 27, the top pressure of the tower is 25 KPa (absolute pressure), the top temperature of the tower is 80 °C, the bottom temperature of the tower is 350 °C, the reflux ratio is 3, and the mass ratio of steam injection at the bottom of the tower to the feed is 0.5.

[0120] The parameters and compositions of each logistics are shown in Table 2.

[0121] Table 2

[0122]

[0123]

[0124] Example 4

[0125] This example provides a method for preparing polyalphaolefin synthetic oil. This method is carried out using the polyalphaolefin synthetic oil preparation system provided in Example 1. This method includes:

[0126] Step 1: The α-olefin raw oil stream A at a temperature of 25 °C and a pressure of 0.05 MPa enters the drying equipment 1 at a flow rate of 1255.2 kg / h for drying and water removal.

[0127] Step 2. The α-olefin raw material oil stream B (with a water content of 10 μg / g) dried by the drying device 1 is mixed with the liquid-phase α-olefin monomer stream J separated by the gas-liquid separation device 7 and fed into the bottom of the polymerization reaction device 2. Meanwhile, the catalyst stream C (pentanol) is fed into the bottom of the polymerization reaction device 2 at a flow rate of 11.2 kg / h. Meanwhile, the gas-phase BF 3 stream D is fed into the top of the polymerization reaction device 2 at a flow rate of 2.1 kg / h, and the α-olefin polymerization reaction is carried out in the polymerization reaction device 2. Among them, the temperature of the polymerization reaction device 2 is 100 °C, the pressure is 2.0 MPaG, and the residence time of the material in the polymerization reaction device 2 is 10 h.

[0128] Step 3. The polymerization reaction product stream E (with a flow rate of about 1404.3 kg / h during the normal operation stage) in the polymerization reaction device 2 enters the stirring and mixing device 3. Meanwhile, the alcohol absorbent stream F (pentanol is used in this embodiment) at a temperature of 25 °C and a pressure of 0.4 MPa (gauge pressure) enters the stirring device 3 at a flow rate of 6.7 kg / h. Meanwhile, the absorbent layer stream G (with a flow rate of about 17.9 kg / h during the normal operation stage) at a temperature of 50 °C and a pressure of 0.8 MPa (gauge pressure) sent out from the separation device 4 is fed into the stirring and mixing device 3 and mixed evenly in the stirring and mixing device 3. Among them, BF 3 forms a complex with the alcohol absorbent to form BF 3 a complex product with alcohol;

[0129] The product mixed evenly in the stirring and mixing device 3 is fed into the separation device 4 for separation to obtain an absorbent layer and a poly-α-olefin synthetic oil layer, and the absorbent layer and the poly-α-olefin synthetic oil layer are respectively fed out from the separation device. Among them, the catalyst system containing BF 3 exists in the absorbent layer, and BF 3 the complex product formed by BF 3 complexing with the alcohol absorbent and the catalyst are dissolved in the alcohol absorbent. Among them, a part of the absorbent layer sent out from the separation device 4 is used as the absorbent layer stream G and fed into the stirring and mixing device 3.

[0130] Step 4: The laminar flow H of poly-α-olefin synthetic oil transported out from the stratifying device 4 is heated to 300 °C by the heating device 5 and then enters the middle section of the monomer stripping tower 6. At the same time, the nitrogen gas stream I enters the monomer stripping tower 6 from the bottom of the monomer stripping tower 6. α-olefin monomer nitrogen stripping separation is carried out in the monomer stripping tower 6. The α-olefin monomer and nitrogen gas are taken out from the top of the monomer stripping tower 6 as a gas phase and enter the gas-liquid separation device 7 for cooling and then gas-liquid separation to obtain the liquid-phase α-olefin monomer stream J. The liquid-phase α-olefin monomer stream J is mixed with the α-olefin raw material oil stream B (water content is 10 μg / g) dried by the drying device 1 for recycling; the poly-α-olefin synthetic oil crude product stream K is output from the bottom of the monomer stripping tower 6.

[0131] Step 5: The poly-α-olefin synthetic oil crude product stream K output from the bottom of the monomer stripping tower 6 is boosted to 6.0 MPa (gauge pressure) by a pump, mixed with the hydrogen gas stream L with a flow rate of 78 kg / h and the hydrogen gas output from the gas phase outlet of the high-pressure separator 9, heated to 300 °C by a heating furnace, and enters the hydrogenation reaction device 8 for hydrogenation reaction. The hydrogenation reaction product in the hydrogenation reaction device 8 is transported to the high-pressure separator 9 for hydrogen separation. The liquid-phase product after hydrogen separation in the high-pressure separator 9 is transported to the low-pressure separator 10 for hydrogen separation. The liquid-phase product after hydrogen separation in the low-pressure separator 10 is transported to the impurity stripping tower 11 for nitrogen stripping to obtain the liquid-phase product, which is the hydrorefined product. Among them, the hydrogen gas output from the gas phase outlet after hydrogen separation in the high-pressure separator 9 is boosted to 5 MPa (gauge pressure) by the recycle hydrogen compressor 12 and then mixed with the hydrogen gas stream L with a flow rate of 78 kg / h and the product of the poly-α-olefin synthetic oil crude product stream K output from the bottom of the monomer stripping tower 6 boosted to 2.0 MPa (gauge pressure); the gas phase output from the gas phase outlet after hydrogen separation in the low-pressure separator 10 is cooled and then discharged to the flare.

[0132] Among them, the temperature of the high-pressure separator 9 is 255 °C and the pressure is 4 MPa (gauge pressure);

[0133] Among them, the overall volume space velocity of the hydrogenation reaction is 0.2 h -1 ;

[0134] Among them, during the nitrogen stripping process, the temperature of the nitrogen gas used is 25 °C, the pressure is 0.4 MPa (gauge pressure), and the mass ratio of gas to oil is 50:1;

[0135] Among them, the pressure of the impurity stripping tower is 0.1 MPa (gauge pressure), the top temperature is 190 °C, the bottom temperature is 220 °C, the number of theoretical plates is 17, and the feed plate position is the top feed of the raw material oil tower.

[0136] Step 6: Heat the hydrofined product output from the impurity stripping column 11 to 370°C in the heating furnace 16 and then transport it to the atmospheric column 13 for atmospheric fractionation; the number of theoretical plates of the atmospheric column 13 is 25, the feed plate position is 18, the top temperature is 105°C, the bottom temperature is 270°C, the reflux ratio is 3, and the mass ratio of bottom steam injection to feed is 1.0. Obtain the PAO base oil product at the top of the atmospheric column 13.

[0137] Heat the bottom material output from the atmospheric column 13 to 380°C in the heating furnace 18 and then transport it to the vacuum column 14 for vacuum fractionation. From the bottom to the top, obtain PAO base oil product stream M, PAO base oil product stream N, PAO base oil product stream O, and PAO base oil product stream P; the number of theoretical plates of the vacuum column 14 is 39, the feed plate position is 27, the top pressure is 50 KPa (absolute pressure), the top temperature is 70°C, the bottom temperature is 350°C, the reflux ratio is 0.1, and the mass ratio of bottom steam injection to feed is 0.1.

[0138] The parameters and compositions of each logistics are shown in Table 3.

[0139] Table 3

[0140]

[0141]

[0142] The above are only the preferred embodiments of the present invention and are not intended to limit the present invention. Any modifications, equivalent replacements, improvements, etc. made within the spirit and principle of the present invention shall be included within the protection scope of the present invention.

Claims

1. A poly-α-olefin synthetic oil preparation system, wherein, the system includes a polymerization reaction device, a stirring and mixing device, a layering device, a monomer separation unit, a hydrofining unit, and a product fractionation unit; the feed inlet of the stirring and mixing device is connected to the discharge outlet of the polymerization reaction device, the discharge outlet of the stirring and mixing device is connected to the feed inlet of the layering device, the absorbent inlet of the stirring and mixing device is respectively connected to the absorbent source and the poly-α-olefin synthetic oil layer outlet of the layering device, the poly-α-olefin synthetic oil layer outlet of the layering device is connected to the feed inlet of the monomer separation unit, the poly-α-olefin synthetic oil outlet of the monomer separation unit is connected to the feed inlet of the hydrofining unit, and the poly-α-olefin synthetic oil outlet of the hydrofining unit is connected to the feed inlet of the product fractionation unit; wherein, a fluid discharge pipeline is provided on the pipeline connecting the absorbent inlet of the stirring and mixing device and the poly-α-olefin synthetic oil layer outlet of the layering device.

2. The system according to claim 1, wherein, The polymerization reaction equipment is provided with a feed inlet, BF 3 Inlet, solid-liquid catalyst inlet, and discharge port; wherein the feed port and solid-liquid catalyst inlet are arranged at the bottom of the polymerization reaction equipment, and there is BF 3 The inlet is arranged at the top of the polymerization reaction device; preferably, a stirrer is provided in the polymerization reaction device; preferably, the polymerization reaction device is provided with a temperature control part, and the temperature control part is provided with a heat transfer medium inlet and a heat transfer medium outlet.

3. The system according to claim 1, wherein, the poly-α-olefin synthetic oil preparation system further includes a drying device, and the discharge outlet of the drying device is connected to the feed inlet of the polymerization reaction device.

4. The system according to claim 1, wherein, the monomer outlet of the monomer separation unit is connected to the feed inlet of the polymerization reaction device.

5. The system according to claim 1, wherein, a sampling point is provided on the pipeline connecting the absorbent inlet of the stirring and mixing device and the poly-α-olefin synthetic oil layer outlet of the layering device.

6. The system according to claim 1, wherein, the monomer separation unit includes a heating device and a monomer stripping tower; the feed inlet of the heating device is connected to the poly-α-olefin synthetic oil layer outlet of the layering device, and the discharge outlet of the heating device is connected to the feed inlet of the monomer stripping tower; preferably, the monomer separation unit further includes a gas-liquid separation device; the gas outlet of the monomer stripping tower is connected to the feed inlet of the gas-liquid separation device, and the liquid outlet of the gas-liquid separation device is used as the monomer outlet of the monomer separation unit.

7. The system according to claim 1, wherein, the hydrofining unit includes a hydrogenation reaction device, a hydrogen separation device, and a stripping tower; the feed inlet of the hydrogenation reaction device is connected to the poly-α-olefin synthetic oil outlet of the monomer separation unit, the discharge outlet of the hydrogenation reaction device is connected to the feed inlet of the hydrogen separation device, the liquid phase outlet of the hydrogen separation device is connected to the feed inlet of the impurity stripping tower, and the poly-α-olefin synthetic oil outlet of the impurity stripping tower is used as the poly-α-olefin synthetic oil outlet of the hydrofining unit; preferably, the hydrogen separation device includes a high-pressure separator and a low-pressure separator; the feed inlet of the high-pressure separator is connected to the discharge outlet of the hydrogenation reaction device, the liquid phase outlet of the high-pressure separator is connected to the feed inlet of the low-pressure separator, and the liquid phase outlet of the low-pressure separator is connected to the feed inlet of the impurity stripping tower; more preferably, the gas phase outlet of the high-pressure separator is connected to the hydrogenation reaction device to realize the hydrogen separated by the high-pressure separator to be returned to the hydrogenation reaction device for recycling; Preferably, the number of theoretical plates of the impurity stripping column is 3 - 25. The outlet for the synthetic polyalphaolefin oil of the impurity stripping column is arranged at the bottom of the impurity stripping column, and the stripping gas outlet of the impurity stripping column is arranged at the top of the impurity stripping column.

8. The system according to claim 1, wherein, the product fractionation unit includes an atmospheric column and a vacuum column. The feed inlet of the atmospheric column is connected to the outlet of the synthetic polyalphaolefin oil of the hydrofining unit. The bottom material outlet of the atmospheric column is connected to the feed inlet of the vacuum column. The side line and bottom material outlets of the vacuum column serve as the outlets for PAO base oil products with different viscosities; Preferably, the number of theoretical plates of the atmospheric column is 20 - 40, and the feed plate position is 15 - 30; Preferably, the number of theoretical plates of the vacuum column is 35 - 65, and the feed plate position is 15 - 30.

9. A method for preparing synthetic polyalphaolefin oil, which uses the synthetic polyalphaolefin oil preparation system according to any one of claims 1 - 8, and this method comprises: Feed the α-olefin feedstock and the catalyst system containing BF 3 to a polymerization reaction apparatus, and conduct the polymerization reaction of the α-olefin feedstock under the catalyst system containing BF 3 in the polymerization reaction apparatus; the catalyst system containing BF 3 comprises gaseous BF 3 and a catalyst; The polymerization reaction product in the polymerization reaction equipment, the alcohol absorbent, and the absorbent layer transported out of the stratification equipment are transported to the stirring and mixing equipment for mixing; among them, BF 3 Complexes with the alcohol absorbent to form BF 3 Complex product with alcohol; The product after mixing evenly in the stirring and mixing equipment is transported to a stratifying equipment for stratification to obtain an absorbent layer and a poly-α-olefin synthetic oil layer, and the absorbent layer and the poly-α-olefin synthetic oil layer are respectively transported out of the stratifying equipment; wherein, a catalyst system containing BF 3 exists in the absorbent layer, and the BF 3 complex formed by BF 3 and the alcohol absorbent and the catalyst are dissolved in the alcohol absorbent; wherein, a part of the absorbent layer transported out of the stratifying equipment is transported to the stirring and mixing equipment; transporting the synthetic polyalphaolefin oil layer transported out from the stratifying device to the monomer separation unit for separating alpha-olefin monomers to obtain a crude synthetic polyalphaolefin oil product; transporting the crude synthetic polyalphaolefin oil product in the monomer separation unit to the hydrofining unit for hydrofining; transporting the hydrofined product in the hydrofining unit to the product fractionation unit for fractionation to obtain PAO base oil products.

10. The method according to claim 9, wherein, Feed the α-olefin feedstock and the catalyst system containing BF 3 to the polymerization reaction equipment, and carry out the polymerization reaction of the α-olefin feedstock under the catalyst system containing BF 3 by the following method: Transport the α-olefin feedstock and the catalyst from the bottom of the polymerization reaction equipment to the polymerization reaction equipment, and at the same time transport gaseous BF 3 from the top of the polymerization reaction equipment to the polymerization reaction equipment for pressure regulation, and carry out the polymerization reaction of the α-olefin feedstock under the catalyst system containing BF 3 in the polymerization reaction equipment.

11. The method according to claim 9, wherein, the catalyst includes at least one of a monohydric alcohol and a dihydric alcohol; preferably, the catalyst includes at least one of a monohydric alcohol with a carbon number not exceeding 10 and a dihydric alcohol with a carbon number not exceeding 10; and / or the alcohol absorbent includes at least one of monohydric alcohols with a carbon number of 1 - 10.

12. The method according to claim 9, wherein, the temperature of the polymerization reaction is 10 - 100 °C, the pressure is 0.05 - 2.0 MPaG, and the time is 0.2 - 40 h.

13. The method according to claim 9, wherein, α-olefin feedstock, BF 3 The mass ratio of the catalyst is: 100:0.0001 - 0.5:0.1 - 1.

14. The method according to claim 9, wherein, The mass ratio of the alcohol absorbent delivered to the stirring and mixing equipment to BF delivered to the polymerization reaction equipment is: 4 - 40:

1. 3 ​ 15. The method according to claim 9, wherein, the alpha-olefin feedstock oil contains at least one of alpha-olefins with 4 - 14 carbon atoms; Preferably, in the alpha-olefin feedstock oil, the mass content of alpha-olefins with 4 - 14 carbon atoms ≥ 80%; Preferably, the mass content of normal alpha-olefins with 4 - 14 carbon atoms in the alpha-olefin feedstock oil ≥ 60%.

16. The method according to claim 9, wherein, transporting the synthetic polyalphaolefin oil layer transported out from the stratifying device to the monomer separation unit for separating alpha-olefin monomers includes: heating the synthetic polyalphaolefin oil layer transported out from the stratifying device to 100 - 300 °C and then transporting it to the monomer stripping column, and carrying out alpha-olefin monomer separation in the monomer stripping column by means of nitrogen stripping; Preferably, the separated gas enters a gas-liquid separation device for cooling and then gas-liquid separation. The separated liquid phase is the alpha-olefin monomer; more preferably, the alpha-olefin monomer is transported to the polymerization reaction device for recycling; Preferably, the temperature of the monomer stripping tower is 10 - 200 °C, the pressure is 5 KPa - 150 KPa, the mass ratio of gas to oil is 0.01 - 0.5, and the injection pressure of nitrogen is 0.10 - 2 MPa.

17. According to the method described in claim 9, wherein, transporting the crude poly-α-olefin synthetic oil product in the monomer separation unit to the hydrofining unit for hydrofining includes: boosting the pressure of the crude poly-α-olefin synthetic oil product output from the monomer separation unit to 2 - 6 MPa, mixing it with hydrogen, heating it to 150 - 300 °C, and performing a hydrogenation reaction in a hydrogenation reaction device; transporting the hydrogenation reaction product in the hydrogenation reaction device to a hydrogen separator for hydrogen separation; transporting the liquid-phase product after hydrogen separation in the hydrogen separator to an impurity stripping tower for nitrogen stripping, and the obtained liquid oxygen product is the hydrofined product; Preferably, the overall volume space velocity of the hydrogenation reaction is 0.1 - 1 h -1 , and the standard volume ratio of hydrogen to oil is 10:1 - 1000:1; Preferably, during the nitrogen stripping process, the temperature of the nitrogen used is 10 - 200 °C, the pressure is 0.10 - 2 MPa, and the mass ratio of gas to oil is 50 - 100:1; Preferably, the tower pressure of the impurity stripping tower is 0 - 1 MPa, the top temperature is 10 - 200 °C, and the bottom temperature is 10 - 200 °C; Preferably, the transporting the hydrogenation reaction product in the hydrogenation reaction device to a hydrogen separator for hydrogen separation includes: transporting the hydrogenation reaction product in the hydrogenation reaction device to a high-pressure separator for hydrogen separation, transporting the liquid-phase product after hydrogen separation in the high-pressure separator to a low-pressure separator for hydrogen separation, and the liquid-phase product after hydrogen separation in the low-pressure separator is used as the liquid-phase product after hydrogen separation in the hydrogen separator; More preferably, the gas-phase product after hydrogen separation in the high-pressure separator is mixed with the product after boosting the pressure of the crude poly-α-olefin synthetic oil product output from the monomer separation unit to 2 - 6 MPa.

18. According to the method described in claim 9, wherein, transporting the hydrofined product in the hydrofining unit to the product fractionation unit for fractionation to obtain a PAO base oil product includes: heating the hydrofined product output from the hydrofining unit to 300 - 380 °C and then transporting it to an atmospheric tower for atmospheric fractionation; heating the bottom material output from the atmospheric tower to 300 - 380 °C and then transporting it to a vacuum tower for vacuum fractionation to obtain a PAO base oil product; Preferably, the top pressure of the atmospheric tower is 0.01 - 0.1 MPa, the top temperature is 200 - 300 °C, the bottom temperature is 250 - 380 °C, the reflux ratio is 0.1 - 20, and the mass ratio of steam injection at the bottom to the feed is 0.01 - 1; Preferably, the top pressure of the vacuum tower is 5 - 50 KPa, the top temperature is 40 - 100 °C, the bottom temperature is 100 - 380 °C, and the mass ratio of steam injection at the bottom to the feed is 0.01 - 1.

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