Piperidine production system and method for continuously producing piperidine

Abstract

The invention relates to the field of piperidine production, and discloses a piperidine production system and a method for continuously producing piperidine. The production system comprises a feeding unit, a vaporization unit, a reaction unit, a cooling unit and a product separation and purification unit which are connected in sequence, the vaporization unit comprises an evaporator, a first gas-liquid separation tank and an electric heater which are connected in sequence; the evaporator sequentially comprises a distribution unit and a heating unit from bottom to top in the axial direction of the evaporator, the distribution unit comprises a gas distribution pipe which is not parallel to the axial direction of the evaporator, through holes are formed in the pipe wall of the gas distribution pipe, a gas phase inlet is formed in one end of the gas distribution pipe, and a liquid phase inlet is formed in the bottom of the evaporator. After passing through the gas distribution pipe, the gas phase is mixed with the liquid phase and enters the heating unit; and a baffle plate is arranged in the first gas-liquid separation tank. The production method has the characteristics of high pyridine conversion rate, high piperidine selectivity and high piperidine purity.

Description

Technical Field

[0001] The invention relates to the field of piperidine production, and in particular to a piperidine production system and a method for continuously producing piperidine. Background Art

[0002] Piperidine, also known as hexahydropyridine, is a colorless liquid with a peppery odor. It is miscible in water and soluble in ethanol, ether, acetone, and benzene. Its alkalinity is slightly stronger than pyridine, and it can form salts with acids. Its chemical properties are similar to those of fatty secondary amines. A strong organic base, it reacts with inorganic acids to form salts. It is an organic synthesis intermediate primarily used in the synthesis of the plant growth regulator mepiquat and the rubber vulcanizer DPTT. It is also used in resin curing agents, local anesthetics, analgesics, fungicides, wetting agents, and other applications. The domestic supply of piperidine primarily relies on imports from the United States, Germany, Japan, the United Kingdom, and India. Therefore, the process of producing piperidine from pyridine hydrogenation is of great significance to the development of my country's DPTT-4 / 6 (dipentamethylthiuram tetra- or hexasulfide) industry.

[0003] Traditional industrial hydrogenation production typically uses an autoclave as the reaction vessel. Solid catalysts and feedstock solutions are added to the autoclave, and the process is performed intermittently under high temperature (200°C) and high pressure (7 MPa). Intermittent hydrogenation requires frequent catalyst separation, making the process cumbersome and posing safety risks. Furthermore, intermittent hydrogenation processes are unsafe and inefficient. Summary of the Invention

[0004] The present invention aims to overcome the problems of low pyridine conversion rate, piperidine selectivity and piperidine purity in the piperidine production process in the prior art, and to provide a piperidine production system and a method for continuously producing piperidine. The production method has the characteristics of high pyridine conversion rate, piperidine selectivity and piperidine purity.

[0005] In order to achieve the above object, the first aspect of the present invention provides a piperidine production system, wherein the production system comprises a feeding unit, a vaporization unit, a reaction unit, a cooling unit and a product separation and purification unit connected in sequence;

[0006] Wherein, the feeding unit comprises a circulating hydrogen buffer tank, a hydrogen compressor and a hydrogen buffer tank connected in sequence;

[0007] Wherein, the vaporization unit comprises an evaporator, a first gas-liquid separation tank and an electric heater connected in sequence;

[0008] From bottom to top along the axial direction of the evaporator, the evaporator includes a distribution unit and a heating unit in sequence. The distribution unit includes a gas distribution pipe arranged non-parallel to the axial direction of the evaporator. The pipe wall of the gas distribution pipe is provided with a through hole. One end of the gas distribution pipe is provided with a gas phase inlet. The bottom of the evaporator is provided with a liquid phase inlet, so that the gas phase passes through the gas distribution pipe, mixes with the liquid phase, and enters the heating unit;

[0009] The first gas-liquid separation tank is provided with a baffle;

[0010] The electric heater is connected to the gas phase outlet of the first gas-liquid separation tank and is used to heat the gas phase flow from the first gas-liquid separation tank into superheated steam;

[0011] Wherein, the reaction unit includes a hydrogenation reactor for hydrogenating hydrogen and pyridine;

[0012] Wherein, the cooling unit comprises a heat exchanger and a cooler connected in sequence, which are optionally used to cool the liquid phase flow from the reaction unit;

[0013] The product separation and purification unit is used to separate and purify the liquid phase flow from the cooling unit to obtain piperidine.

[0014] Preferably, the through hole has a diameter of 6-30 mm, preferably 8-18 mm.

[0015] Preferably, the number of the through holes is 9-50, preferably 20-45.

[0016] Preferably, the distance between any two adjacent through holes along the axial direction of the gas distribution pipe is 10-36 mm, preferably 15-30 mm.

[0017] Preferably, the first gas-liquid separation tank is provided with 1-5 levels of baffles, preferably 2-4 levels of baffles.

[0018] A second aspect of the present invention provides a method for continuously producing piperidine using the production system described in the first aspect, wherein the method comprises:

[0019] (1) The hydrogen is passed into the circulating hydrogen buffer tank, pressurized by the hydrogen compressor, and then enters the hydrogen buffer tank;

[0020] (2) introducing hydrogen and pyridine in the hydrogen buffer tank into a raw material vaporization device for vaporization to obtain a vaporized raw material;

[0021] The vaporization comprises introducing hydrogen from the gas phase inlet (N4) and pyridine from the liquid phase inlet (N5) at the bottom of the evaporator, so that the hydrogen passes through the gas distribution pipe, mixes with the pyridine, and enters the heating unit for a first heating; sending the heated logistics into a first gas-liquid separation tank for gas-liquid separation; and sending the gas phase logistics obtained by the gas-liquid separation into an electric heater for a second heating;

[0022] (3) passing the vaporized feedstock into a hydrogenation reactor, and performing a hydrogenation reaction in the presence of a catalyst and optionally a diluent to obtain a hydrogenated product;

[0023] (4) Cooling and separating the hydrogenation product.

[0024] Preferably, the molar ratio of hydrogen to pyridine is 8-40:1, more preferably 10-30:1.

[0025] Through the above technical solution, compared with the prior art, the present invention has the following beneficial effects:

[0026] (1) Using solvent-free pyridine as raw material reduces the use of solvent and subsequent solvent recovery, thus reducing process complexity and cost;

[0027] (2) The raw material vaporization device can completely vaporize pyridine, thereby improving the pyridine conversion rate, piperidine selectivity and purity. BRIEF DESCRIPTION OF THE DRAWINGS

[0028] Figure 1 A flow chart for the piperidine production system and apparatus;

[0029] Figure 2 This is a schematic diagram of the evaporation device;

[0030] Figure 3 It is the structural diagram of the evaporator and the partial enlarged diagram of point C;

[0031] Figure 4 for Figure 3 A front view of the section at point I in the evaporator and cross-sectional views along lines AA and BB;

[0032] Figure 5 Schematic diagram of the structure of the first gas-liquid separation tank.

[0033] Description of Reference Numerals

[0034] N1-discharge port; N2-heating medium inlet; N3-heating medium outlet; N4-gas phase inlet; N5-liquid phase inlet; N6-gas phase discharge port; N7-first gas-liquid separation tank inlet; N8-liquid phase discharge port; L1-on-site liquid level gauge; L2-on-site liquid level gauge; P-on-site pressure gauge; T-on-site thermometer; H-hand hole. DETAILED DESCRIPTION

[0035] The endpoints of the ranges and any values disclosed herein are not limited to the precise ranges or values, and these ranges or values should be understood to include values close to these ranges or values. For numerical ranges, the endpoints of each range, the endpoints of each range and individual point values, and the individual point values can be combined with each other to obtain one or more new numerical ranges, which should be considered to be specifically disclosed herein.

[0036] In the present invention, unless otherwise specified, directional terms such as "upper, lower, left, right" generally refer to directions shown in the accompanying drawings. "Inside" and "outside" refer to inside and outside relative to the outline of each component itself.

[0037] In the present invention, unless otherwise specified, the "bottom" or "lower end" of the reactor refers to 90-100% of the length of the reactor from top to bottom, and the "upper end" of the reactor refers to 0-10% of the length of the reactor from top to bottom.

[0038] In the present invention, unless otherwise specified, the terms "first" and "second" do not limit the materials, components, or steps, but are used only to distinguish them. For example, the "first" and "second" in "first gas-liquid separator" and "second gas-liquid separator" simply indicate that they are not the same gas-liquid separator. Similarly, "first heating" and "second heating" simply indicate that they are not the same heating.

[0039] In the present invention, unless otherwise specified, components and units are connected by pipelines.

[0040] A first aspect of the present invention provides a piperidine production system, wherein the production system comprises a feeding unit, a vaporization unit, a reaction unit, a cooling unit and a product separation and purification unit connected in sequence;

[0041] Wherein, the feeding unit comprises a circulating hydrogen buffer tank, a hydrogen compressor and a hydrogen buffer tank connected in sequence;

[0042] Wherein, the vaporization unit comprises an evaporator, a first gas-liquid separation tank and an electric heater connected in sequence;

[0043] From bottom to top along the axial direction of the evaporator, the evaporator includes a distribution unit and a heating unit in sequence. The distribution unit includes a gas distribution pipe arranged non-parallel to the axial direction of the evaporator. The pipe wall of the gas distribution pipe is provided with a through hole. One end of the gas distribution pipe is provided with a gas phase inlet. The bottom of the evaporator is provided with a liquid phase inlet, so that the gas phase passes through the gas distribution pipe, mixes with the liquid phase, and enters the heating unit;

[0044] The first gas-liquid separation tank is provided with a baffle;

[0045] The electric heater is connected to the gas phase outlet of the first gas-liquid separation tank and is used to heat the gas phase flow from the first gas-liquid separation tank into superheated steam;

[0046] Wherein, the reaction unit includes a hydrogenation reactor for hydrogenating hydrogen and pyridine;

[0047] Wherein, the cooling unit comprises a heat exchanger and a cooler connected in sequence, which are optionally used to cool the liquid phase flow from the reaction unit;

[0048] The product separation and purification unit is used to separate and purify the liquid phase flow from the cooling unit to obtain piperidine.

[0049] like Figure 3 As shown, the gas phase inlet of the gas distribution pipe is located outside the evaporator, and the gas distribution pipe is located inside.

[0050] The present invention improves the pyridine conversion rate, piperidine selectivity and piperidine purity by sequentially connecting a feeding unit, a vaporization unit, a reaction unit, a cooling unit and a product separation and purification unit. The evaporator in the evaporation device of the vaporization unit is equipped with a gas distribution pipe with a through hole, so that the gas phase can be evenly distributed in a sufficiently short time. After the logistics is heated in the evaporator, it enters the first gas-liquid separation tank. The baffle provided in the first gas-liquid separation tank guides the pyridine entering the first gas-liquid separation tank, firstly, to fully separate the gas and liquid, and secondly, to keep the liquid level in the first gas-liquid separation tank stable. The electric heater performs the final superheated steam conversion on the separated gas phase, so that the entire system can stably and efficiently achieve the goal of complete vaporization of the gas and liquid phases.

[0051] In the present invention, "optionally" means that a heat exchanger and a cooler can be used to cool the hydrogenated product from the hydrogenation reactor, or only a cooler can be used to cool the hydrogenated product from the hydrogenation reactor. According to a specific embodiment of the present invention, preferably, the present invention uses a heat exchanger and a cooler together to cool the hydrogenated product from the hydrogenation reactor.

[0052] In the present invention, hydrogen is exchanged with the hydrogenation product from the hydrogenation reactor in the heat exchanger, which can save energy consumption, help reduce production costs and better cool the hydrogenation product.

[0053] In the present invention, superheated steam refers to a gaseous mixture obtained by heating with an electric heater. For example, when the feed gas phase is hydrogen and the feed liquid phase is pyridine, the superheated steam is a gaseous mixture of hydrogen and pyridine.

[0054] In the present invention, the axial direction refers to the center line direction determined by the length direction of the evaporator itself. The non-parallel arrangement means that the axial direction of the gas distribution pipe forms a certain angle with the axial direction of the evaporator.

[0055] The present invention does not particularly limit the angle between the gas distribution pipe and the axial direction of the evaporator, as long as it can satisfy the passage of the gas phase. Those skilled in the art can select it according to actual needs. Preferably, the gas distribution pipe is perpendicular to the axial direction of the evaporator.

[0056] Preferably, the distribution unit of the evaporator includes a gas distribution pipe and a support beam for fixing the gas distribution pipe.

[0057] The present invention does not particularly limit the arrangement of the support beam, as long as it can fix the gas distribution pipe without shaking. Those skilled in the art can choose according to actual needs.

[0058] According to the present invention, preferably, along the axial direction of the gas distribution pipe, there is a gap between the gas distribution pipe and the right side tube wall of the evaporator for liquid phase circulation.

[0059] According to the present invention, preferably, the gas phase stream is introduced from the gas phase inlet N4.

[0060] The present invention has a wide range of selection for the diameter of the gas distribution pipe. In order to further meet the required flow rate of the gas phase logistics, preferably, the diameter of the gas distribution pipe is 0.1-0.8m, more preferably 0.18-0.5m.

[0061] The present invention does not particularly limit the distance between the gas distribution pipe and the bottom of the evaporator, as long as the gas phase logistics can fully enter the heating unit. Those skilled in the art can make a choice according to actual needs. Preferably, the distance between the gas distribution pipe and the bottom of the evaporator is 0.05-0.6m, more preferably 0.1-0.4m.

[0062] The present invention has a wide range of selection for the number of groups of through holes on the circumferential direction of the tube wall of the gas distribution tube. In order to further enable the gas phase to be evenly distributed in a sufficiently short time, preferably, n groups of through holes are provided along the circumference of the tube wall of the gas distribution tube, and each group independently includes m through holes arranged along the axial direction, n≥3, m≥3, for example, it can be n=3, m=3, n=3, m=4, n=3, m=5, n=4, m=5, n=3, m=8, n=4, m=10, n=4, and any value in the range of any two combinations of n and m; two adjacent groups of through holes are preferably staggered.

[0063] According to the present invention, preferably, Figure 4 As shown, no through hole is provided on the pipe wall above the horizontal plane where the axis of the gas distribution pipe is located (excluding the horizontal plane). It should be noted that in the present invention, the position of the through hole refers to the position of the center of the through hole (if it is a circular through hole, it is the center of the circle).

[0064] The present invention has a wide range of selection for the diameter of the through hole. In order to further enable the gas phase to be evenly distributed in a sufficiently short time, preferably, the diameter of the through hole is 6-30 mm, more preferably 8-18 mm.

[0065] The aperture of the through hole in the present invention is within the above preferred range, which is more conducive to the uniform distribution of gas phase flow.

[0066] The present invention has a wide range for selecting the number of through holes. In order to further enable the gas phase to be evenly distributed in a sufficiently short time, preferably, the number of the through holes is 9-50, and more preferably 20-45.

[0067] In the present invention, the number of through holes is within the above preferred range, which is more conducive to enhancing the dispersion effect of the gas phase flow.

[0068] The present invention has a wide range of selection for the spacing between any two adjacent through holes along the axial direction of the gas distribution pipe. In order to further enable the gas phase to be evenly distributed in a sufficiently short time, preferably, the spacing between any two adjacent through holes along the axial direction of the gas distribution pipe is 10-36mm, more preferably 15-30mm.

[0069] It should be noted that the spacing between two adjacent through holes in the present invention refers to the distance between the centers of the two holes.

[0070] The preferred range of the distance between two adjacent holes in the through-holes of the present invention is more conducive to uniform distribution of gaseous flow, so that the flow is in a completely vaporized state before entering the hydrogenation reactor.

[0071] The present invention has a wide range of selection for the diameter ratio of the axial length of the evaporator and the gas distribution pipe. In order to further enable the gas phase to be evenly distributed in a sufficiently short time, preferably, the diameter ratio of the axial length of the evaporator and the gas distribution pipe is 1:0.5-0.9, for example, it can be 1:0.5, 1:0.55, 1:0.6, 1:0.7, 1:0.75, 1:0.8, 1:0.85, 1:0.9 and any value in the range of any two combinations, more preferably 1:0.6-0.9.

[0072] The present invention does not particularly limit the shape of the through hole, as long as the gas phase can be evenly distributed in a sufficiently short time. Those skilled in the art can choose according to actual needs. Preferably, the shape of the through hole is selected from at least one of a circular hole, a triangular hole, a square hole, a polygonal hole, an elliptical hole, and a diamond hole, and a circular hole is more preferred.

[0073] The gas distribution pipe of the present invention has a wide range of settings. The preferred embodiments do not limit the scope of the present invention, but are only used to explain the solutions in the embodiments of the present invention. Those skilled in the art can choose according to actual needs. According to a preferred embodiment of the present invention, Figure 4 As shown, the axial length-to-diameter ratio of the evaporator and the gas distribution pipe is 1:0.6-0.9. The gas distribution pipe includes three groups of through holes, totaling 20-45 through holes. Each group of through holes includes at least three through holes arranged axially, and each group of through holes is independent of each other. Two groups of through holes are symmetrically arranged on the pipe wall on a horizontal plane where the axis of the gas distribution pipe lies, while the other group of through holes is arranged on the pipe wall below this horizontal plane and staggered with the two groups of through holes above. Specifically, when viewed from above, the through holes below the horizontal plane are located between two adjacent through holes on the horizontal plane. This is more conducive to extending the residence time of the gas phase in the gas distribution pipe, improving pyridine conversion, piperidine selectivity, and piperidine purity.

[0074] According to the present invention, preferably, the heating unit is arranged 0.1-1.2 m above the gas distribution pipe, more preferably 0.2-0.8 m above the gas distribution pipe.

[0075] The present invention does not particularly limit the number of heat exchange tubes in the heating unit, as long as sufficient vaporization of the gas and liquid phases is achieved. Persons skilled in the art may select a suitable number based on practical needs. Preferably, the heating unit includes at least one set of heat exchange tubes for circulating the heating medium. According to a preferred embodiment of the present invention, the heating unit more preferably comprises two sets of heat exchange tubes (one set comprising 13 heat exchange tubes).

[0076] The present invention does not particularly limit the material, structure, shape and other equipment features of the heat exchange tubes, as long as they can fully vaporize the gas and liquid phase logistics. Those skilled in the art can choose according to actual needs.

[0077] According to the present invention, preferably, the heating medium of the heating unit is water vapor.

[0078] In the present invention, a flange is provided between the heating medium outlet N3 and the gas phase inlet N4 of the evaporator, near the heating medium outlet N3. The shell side is sealed by the flange, so that the logistics enters from the tube side and the heating medium circulates in the shell side.

[0079] According to the present invention, preferably, a heating medium inlet N2 is provided at the upper portion of the evaporator, and a heating medium outlet N3 is provided at the lower portion, so that the heating medium enters the evaporator from the heating medium inlet N2 and flows out from the heating medium outlet N3.

[0080] The present invention does not particularly limit the number of stages of the baffles of the first gas-liquid separation tank. As long as the liquid phase logistics can pass through the baffles in sequence and enter the tank to ensure the stability of the liquid level in the first gas-liquid separation tank, those skilled in the art can make a choice according to actual needs. Preferably, the first gas-liquid separation tank is provided with 1-5 stages of baffles, more preferably 2-4 stages of baffles.

[0081] The present invention does not particularly limit the thickness of the baffle of the first gas-liquid separation tank. As long as the liquid phase logistics can pass through the baffle in sequence and enter the tank, and the liquid level in the first gas-liquid separation tank is guaranteed to be stable, those skilled in the art can choose according to actual needs. Preferably, the thickness of the baffle is 2-10 cm, and more preferably 3-8 cm.

[0082] The present invention does not particularly limit the length of the baffle of the first gas-liquid separation tank. As long as the liquid phase logistics can pass through the baffle in sequence and enter the tank to ensure the stability of the liquid level in the first gas-liquid separation tank, those skilled in the art can choose according to actual needs. Preferably, the length of the baffle is 0.1-0.6m, more preferably 0.2-0.5m.

[0083] The present invention does not particularly limit the bending angle of the baffle of the first gas-liquid separation tank. As long as the liquid phase logistics can pass through the baffle in sequence and enter the tank to ensure the stability of the liquid level in the first gas-liquid separation tank, those skilled in the art can make a selection according to actual needs. Preferably, the bending angle of the baffle is 10-50°, more preferably 10-30°.

[0084] It should be noted that the bending angle has a conventional definition in this field. The baffle bending angle in the present invention refers to the angle formed by the reverse bending of the end of the baffle toward the flow direction of the logistics, which is used to adjust the flow direction of the fluid.

[0085] The present invention does not particularly limit the distance between the baffle and the liquid phase discharge port N8 at the lower end of the first gas-liquid separation tank. As long as the liquid phase logistics can pass through the baffle in sequence and enter the tank to ensure the stability of the liquid level in the first gas-liquid separation tank, those skilled in the art can make a choice according to actual needs. Preferably, the distance between the baffle and the liquid phase discharge port N8 at the lower end of the first gas-liquid separation tank is 0.6-2m, more preferably 0.8-1.4m.

[0086] It should be noted that the distance between the baffle and the liquid phase discharge port N8 at the lower end of the first gas-liquid separation tank in the present invention refers to the vertical distance from the bottom of the last baffle to the liquid phase discharge port.

[0087] The present invention is more conducive to ensuring the stability of the liquid level in the first gas-liquid separation tank by setting the baffle parameters within a preferred range. The stable liquid level is conducive to the accuracy of the on-site liquid level meter measurement, and the liquid will not be re-entrained by the gas phase due to liquid level fluctuations. The baffle acts as a physical barrier in the first gas-liquid separation tank. When the gas-liquid mixture enters the first gas-liquid separation tank, it first hits the baffle. The baffle changes the flow direction of the gas-liquid mixture, causing the liquid phase to flow downward along the baffle surface under the action of gravity, while the gas phase bypasses the baffle and flows upward, which is more conducive to achieving the initial separation of gas and liquid.

[0088] The present invention does not particularly limit the structure of the baffle of the first gas-liquid separation tank. As long as the liquid phase flow can pass through the baffle in sequence and enter the tank to ensure the stability of the liquid level in the first gas-liquid separation tank, technical personnel in this field can make a choice according to actual needs. Preferably, the surface structure of the baffle is serrated.

[0089] It should be noted that, according to a specific embodiment of the present invention, preferably, the specific arrangement of the saw teeth on the baffle of the present invention is that the saw teeth on the baffle are arranged in a dense shape, each saw tooth is small in size, and there is a uniform interval between adjacent saw teeth.

[0090] The present invention does not particularly limit the tooth height of the saw teeth of the baffle of the first gas-liquid separation tank. As long as the liquid phase logistics can pass through the baffle into the tank in sequence and ensure the stability of the liquid level in the first gas-liquid separation tank, those skilled in the art can make a selection according to actual needs. Preferably, the tooth height of the saw teeth is 1-5 cm, more preferably 1-3 cm.

[0091] In the present invention, the tooth height within the preferred range is more conducive to forming a more tortuous flow path for the liquid phase on the serrations, prolonging the residence time of the liquid phase on the baffle, and allowing small droplets in the liquid phase more time to disperse into a thin layer.

[0092] It should be noted that the tooth height of the saw teeth of the baffle in the present invention is the vertical distance from the bottom of the saw teeth (that is, the part connected to the baffle body) to the top of the saw teeth.

[0093] The present invention does not particularly limit the pitch of the saw teeth of the baffle of the first gas-liquid separation tank. As long as the liquid phase logistics can pass through the baffle into the tank in sequence and ensure the stability of the liquid level in the first gas-liquid separation tank, those skilled in the art can make a selection according to actual needs. Preferably, the pitch of the saw teeth is 1-5 cm, more preferably 1-3 cm.

[0094] In the present invention, the tooth pitch within the preferred range is more conducive to evenly distributing the liquid phase between the teeth, avoiding local accumulation of the liquid phase, while also ensuring that the gas phase has a suitable channel to pass smoothly, reducing the interference of the gas phase on the liquid phase.

[0095] It should be noted that the pitch of the baffle teeth in the present invention refers to the distance between the same points on two adjacent teeth. If each tooth is considered an independent geometric shape (such as a triangle), the pitch is the distance from a specific point on one tooth (such as the tip or valley) to the corresponding point on the adjacent tooth.

[0096] According to the present invention, preferably, the first gas-liquid separation tank is provided with an on-site liquid level meter L1 and an on-site liquid level meter L2.

[0097] The present invention does not particularly limit the type of on-site liquid level meter, as long as it can achieve the purpose of testing the liquid level, and those skilled in the art can choose it according to actual needs.

[0098] The present invention does not specifically limit the height difference between the on-site liquid level gauge L1 and the on-site liquid level gauge L2. As long as the two liquid level gauges can accurately measure the liquid level heights at different positions in the first gas-liquid separation tank, those skilled in the art can make a selection according to actual needs. Preferably, the height difference between the on-site liquid level gauge L1 and the on-site liquid level gauge L2 is 0.4-1.4m, preferably 0.6-1.2m.

[0099] In the present invention, the on-site liquid level gauge L1 is located above the on-site liquid level gauge L2 and between the first gas-liquid separation tank feed port N7 and the hand hole H.

[0100] According to the present invention, preferably, the first gas-liquid separation tank is provided with an on-site pressure gauge P.

[0101] The present invention does not particularly limit the type of on-site pressure gauge, as long as it can accurately measure the pressure in the first gas-liquid separation tank. Those skilled in the art can select it according to actual needs.

[0102] The present invention does not particularly limit the distance at which the on-site pressure gauge is set above the first gas-liquid separation tank inlet N7, as long as it can accurately measure the pressure in the first gas-liquid separation tank. Those skilled in the art can make a selection according to actual needs. Preferably, the on-site pressure gauge P is set 0.3-1.8m above the first gas-liquid separation tank inlet N7, and more preferably 0.5-1.4m.

[0103] According to the present invention, preferably, the first gas-liquid separation tank is provided with an on-site thermometer T.

[0104] The present invention does not particularly limit the type of on-site thermometer, as long as it can accurately measure the temperature in the first gas-liquid separation tank. Those skilled in the art can select it according to actual needs.

[0105] The present invention does not particularly limit the distance at which the field thermometer T is set above the field liquid level gauge L1. As long as the temperature in the first gas-liquid separation tank can be accurately measured, those skilled in the art can make a selection according to actual needs. Preferably, the field thermometer T is set 0.5-2m above the field liquid level gauge L1, and more preferably 0.8-1.6m.

[0106] According to the present invention, preferably, the first gas-liquid separation tank is provided with a hand hole H.

[0107] It should be noted that the provision of a hand hole in the present invention facilitates operators to operate the equipment.

[0108] The present invention has no special limitation on the distance at which the hand hole H is set below the first gas-liquid separation tank inlet N7, as long as it can meet the operator's operation needs. Those skilled in the art can choose according to actual needs. Preferably, the hand hole H is set 0.8-2.5m below the first gas-liquid separation tank inlet N7, more preferably 1-2m.

[0109] According to the present invention, preferably, the production system further comprises a preheating unit connected to the raw material vaporization unit.

[0110] According to the present invention, preferably, the preheating unit includes a pyridine preheater.

[0111] The preheating unit in the present invention is more conducive to improving the pyridine conversion rate and piperidine selectivity.

[0112] According to the present invention, preferably, the gas phase outlet of the second gas-liquid separation tank is connected to a circulating hydrogen buffer tank.

[0113] In the present invention, the gas phase outlet of the second gas-liquid separation tank is connected to the circulating hydrogen buffer tank so that the remaining hydrogen can be recycled, which significantly reduces the waste of hydrogen and improves the resource utilization efficiency.

[0114] According to the present invention, preferably, the reaction unit comprises two hydrogenation reactors, and the hydrogenation reactors are shell-and-tube fixed-bed reactors.

[0115] According to the present invention, preferably, the two hydrogenation reactors are connected in series.

[0116] The present invention does not particularly limit the type of heat exchange medium in the shell-and-tube fixed bed reactor, as long as it can meet the temperature of the hydrogenation reaction. Those skilled in the art can select it according to actual needs, for example, it can be heat transfer oil.

[0117] The present invention has no particular limitation on the flow rate of the heat exchange medium, as long as the flow rate of the heat exchange medium can be adjusted to control the temperature of the reactor. Those skilled in the art can select the heat exchange medium according to actual needs.

[0118] Two tubular fixed-bed reactors are used in series, using thermal oil as the heat exchange medium. The reactor temperature is precisely controlled by adjusting the flow of the thermal oil, which can ensure the stability and efficiency of the reaction process.

[0119] According to the present invention, preferably, the product separation and purification unit comprises a second gas-liquid separation tank, a flash separation tank and a distillation tower connected in sequence.

[0120] The use of a distillation tower in the present invention is more conducive to improving the selectivity of pyridine in the pyridine hydrogenation reaction.

[0121] The present invention does not particularly limit the specific arrangement of the second gas-liquid separation tank, as long as it can be used to separate piperidine. Those skilled in the art can select it according to actual needs.

[0122] The present invention does not particularly limit the specific arrangement of the flash separation tank, as long as it can be used to remove residual light components. Those skilled in the art can select it according to actual needs.

[0123] The present invention does not particularly limit the specific arrangement of the distillation tower, as long as it can be used for refining and purification, and those skilled in the art can select it according to actual needs.

[0124] The present invention is more conducive to improving the pyridine conversion rate, piperidine selectivity and piperidine purity by separating in the second gas-liquid separation tank, the flash separation tank and the distillation tower which are connected in sequence.

[0125] A second aspect of the present invention provides a method for continuously producing piperidine using the production system described in the first aspect, wherein the method comprises:

[0126] (1) The hydrogen is passed into the circulating hydrogen buffer tank, pressurized by the hydrogen compressor, and then enters the hydrogen buffer tank;

[0127] (2) introducing hydrogen and pyridine in the hydrogen buffer tank into a raw material vaporization device for vaporization to obtain a vaporized raw material;

[0128] The vaporization includes introducing hydrogen from the gas phase inlet N4 and pyridine from the liquid phase inlet N5 at the bottom of the evaporator, so that the hydrogen passes through the gas distribution pipe, mixes with the pyridine, and enters the heating unit for the first heating; the heated logistics is sent to the first gas-liquid separation tank for gas-liquid separation; the gas phase logistics obtained by gas-liquid separation is sent to the electric heater for the second heating;

[0129] (3) passing the vaporized feedstock into a hydrogenation reactor, and performing a hydrogenation reaction in the presence of a catalyst and optionally a diluent to obtain a hydrogenated product;

[0130] (4) Cooling and separating the hydrogenation product.

[0131] In the present invention, "optionally" means that the catalyst and the diluent can be added to the hydrogenation reactor together, or only the catalyst can be added. Preferably, in the present invention, the catalyst and the diluent are added to the hydrogenation reactor together.

[0132] The use of a diluent in the present invention is more conducive to saving catalyst costs.

[0133] According to the present invention, preferably, the molar ratio of hydrogen to pyridine is 8-40:1, more preferably 10-30:1.

[0134] The present invention adopts the above-mentioned low molar ratio of hydrogen to pyridine, the amount of hydrogen used is small, the molar ratio of hydrogen to pyridine is relatively low, the hydrogen consumption is small, and the remaining hydrogen can be recycled, which significantly reduces the waste of hydrogen and improves the resource utilization efficiency.

[0135] According to the present invention, preferably, the mass ratio of the catalyst to the diluent is 1:1-10, more preferably 1:2-5.

[0136] According to the present invention, preferably, the catalyst includes a metal element and a carrier.

[0137] According to the present invention, preferably, based on the total amount of the catalyst, the content of the metal element is 1-2 wt%, and the content of the carrier is 98-99 wt%.

[0138] According to the present invention, preferably, the metal element is selected from at least one of platinum, rhodium, palladium, nickel, copper, zinc, ruthenium and molybdenum, more preferably at least one of palladium, nickel and ruthenium.

[0139] According to the present invention, preferably, the carrier is selected from at least one of alumina, silica, activated carbon and molecular sieve, more preferably alumina.

[0140] According to the present invention, preferably, the average particle size of the catalyst is 3-5 mm.

[0141] The present invention has no particular limitation on the shape of the catalyst, as long as it can catalyze the hydrogenation reaction. Those skilled in the art can select the catalyst according to actual needs. Preferably, the catalyst is spherical.

[0142] The present invention does not particularly limit the type of diluent, as long as it can save catalyst costs and does not deteriorate the hydrogenation reaction effect. Those skilled in the art can choose according to actual needs. Preferably, the diluent is selected from at least one of quartz sand, alumina ceramic balls, silicon oxide ceramic balls and metal ceramic balls.

[0143] It should be noted that the present invention does not particularly limit the form of the diluent, and those skilled in the art can select it according to actual needs. According to a specific embodiment of the present invention, the diluent is preferably spherical.

[0144] The present invention does not particularly limit the equipment for separation in step (2), as long as it can meet the requirements of the present invention, and those skilled in the art can select it according to actual needs. Preferably, the separation in step (4) is carried out in the second gas-liquid separation tank, the flash separation tank and the distillation tower.

[0145] According to the present invention, preferably, the hydrogen obtained from the second gas-liquid separation tank is returned to the circulating hydrogen buffer tank.

[0146] According to the present invention, preferably, the step (2) further comprises preheating the pyridine in a pyridine preheater.

[0147] According to the present invention, preferably, the pressurizing pressure in step (1) is 1-4 MPa, more preferably 1.5-3.6 MPa.

[0148] The operating pressure of the hydrogenation reactor in the present invention is relatively low, which helps to reduce equipment and operating costs while improving safety.

[0149] According to the present invention, preferably, the first heating conditions include: a temperature of 100-180° C., more preferably 100-150° C.; and a pressure of 1-5 MPa, more preferably 1.5-4 MPa.

[0150] According to the present invention, preferably, the second heating conditions include: a temperature of 100-200° C., more preferably 130-180° C.; and a pressure of 1-5 MPa, more preferably 1.5-4 MPa.

[0151] According to the present invention, preferably, the conditions for hydrogenation in step (3) include: temperature of 120-180°C, pressure of 1-4 MPa, pyridine mass space velocity of 0.1-2 h -1 .

[0152] It should be noted that the pyridine mass space velocity in the present invention refers to the liquid phase feed mass flow rate of pyridine / catalyst loading mass.

[0153] According to the present invention, preferably, the cooling temperature is 0-20°C, more preferably 0-10°C.

[0154] According to the present invention, preferably, in the second gas-liquid separation tank of the product separation and purification unit, the operating temperature is 0-30°C, more preferably 0-10°C, and the operating pressure is 1-4 MPa, more preferably 1.5-3 MPa.

[0155] According to the present invention, preferably, in the flash separation tank of the product separation and purification unit, the flash temperature is 0-50° C., more preferably 0-20° C.; the flash pressure is 0-3 MPa, more preferably 0.1-1 MPa.

[0156] According to the present invention, preferably, in the distillation tower of the product separation and purification unit, the number of plates is 10-50, more preferably 10-30; and the distillation pressure is 0-2 MPa, more preferably 0.05-1 MPa.

[0157] According to some specific embodiments of the present invention, Figure 1 、 Figure 2 、 Figure 3 、 Figure 4 and Figure 5As shown, fresh hydrogen is introduced into the circulating hydrogen buffer tank and mixed with recycled hydrogen from the second gas-liquid separator. The mixed gas then enters the hydrogen compressor for pressurization before entering the hydrogen buffer tank. After pyridine is heated to a certain temperature in the pyridine preheater, the hydrogen in the hydrogen buffer tank and the pyridine in the pyridine preheater are added to the raw material vaporization unit according to the molar ratio of hydrogen to pyridine. In the raw material vaporization unit, low-pressure steam is used as the heating medium. It flows into the evaporator through the heating medium inlet N2. After being heated by the heating unit inside the evaporator, it condenses and flows out of the heating medium outlet N3, thus achieving the purpose of circulating heating. Hydrogen passes through the evaporator's gas phase inlet N4. The evenly distributed through-holes on the distribution pipe ensure that the hydrogen is evenly distributed in a very short time. Pyridine enters the evaporator through the liquid phase inlet N5 at the bottom of the evaporator and mixes with the hydrogen through the gap between the gas distribution pipe and the evaporator. Hydrogen and pyridine form a gas-liquid mixture in the evaporator heating unit, flowing out of the outlet N1 at the top of the evaporator and then entering the first gas-liquid separator tank through the first gas-liquid separator tank inlet N7. The first gas-liquid separator tank is equipped with multiple baffles, allowing the liquid phase to flow into the tank through the baffles in sequence, effectively maintaining a stable liquid level in the tank. The pyridine in the first gas-liquid separator tank is discharged from the liquid phase outlet N8 at the bottom of the first gas-liquid separator tank and then flows back into the evaporator through the liquid phase inlet N5 at the bottom of the evaporator. The gas phase enters the electric heater from the gas phase outlet N6 of the first gas-liquid separator tank, where it is further heated and converted into superheated steam. The gas outlet of the first gas-liquid separator tank is equipped with an on-site thermometer T. This temperature measurement port establishes an interlocking mechanism with the flow rate of the evaporator's heating medium, the water vapor. The flow rate of the evaporator water vapor can be adjusted according to the outlet temperature to ensure the stable operation of the entire system. In addition, the first gas-liquid separation tank is also equipped with an on-site liquid level gauge L1, an on-site liquid level gauge L2, an on-site pressure gauge P, an on-site thermometer T, and a hand hole H. These monitoring and operating components provide convenience and guarantee for the operation monitoring, maintenance and overhaul of the device. After the hydrogen and pyridine are vaporized, the hydrogenation reaction is carried out in the hydrogenation reactor. The hydrogenation temperature is controlled by heat transfer oil. The hydrogenation reaction is catalyzed by the catalyst and diluent. In the heat exchanger, the hydrogen and the hydrogenation product are heat-exchanged to reduce the temperature of the hydrogenation product and save energy. The hydrogenation product after heat exchange in the heat exchanger is cooled in a cooler and enters the second gas-liquid separation tank for separation. The gas phase enters the circulating hydrogen buffer tank, and the liquid phase enters the flash separation tank for further separation. The gas that cannot be liquefied in the flash separation tank is discharged, and the liquid phase enters the distillation tower for refining and purification. The product piperidine is obtained at the top of the tower, and the heavy components are discharged from the bottom of the tower.

[0158] After the above process engineering, the production process of piperidine by hydrogenation of hydrogen and pyridine is simple, and can achieve complete vaporization of pyridine. It has the characteristics of high pyridine conversion rate, piperidine selectivity and piperidine purity. The raw material pyridine does not require a solvent, which reduces the subsequent solvent removal process. The hydrogen can be recycled, reducing hydrogen waste.

[0159] The present invention will be described in detail below through examples.

[0160] In the following examples and comparative examples, unless otherwise specified, conventional methods are used; the reagents and materials used, unless otherwise specified, are commercially available and / or prepared using methods known in the art.

[0161] In the following examples and comparative examples, the catalysts are spherical in shape.

[0162] In the following examples and comparative examples, The conversion of pyridine was determined by gas chromatography.

[0163] In the following examples and comparative examples, The selectivity of piperidine was determined by gas chromatography.

[0164] In the following examples and comparative examples, The purity of piperidine was determined by gas chromatography.

[0165] Example 1

[0166] The production system used is Figure 1 As shown, the evaporation device structure is as follows Figure 2 As shown, the axial length to diameter ratio of the evaporator and the gas distribution pipe is 1:0.88, the gas distribution pipe comprises three groups of through holes, 30 evenly distributed circular holes are provided, each group of through holes comprises 10 through holes arranged along the axial direction, and each group of through holes is independent of each other. Figure 4 Similar to the structure shown, two groups of through holes are symmetrically arranged on the tube wall on the horizontal plane where the axis of the gas distribution pipe is located, and another group of through holes is arranged on the tube wall below the horizontal plane, and is staggered with the two groups of through holes above. Specifically, when viewed from above, the through holes below the horizontal plane are located between two adjacent through holes on the horizontal plane. The aperture of the through hole is 12 mm, and the distance between any two adjacent through holes along the axial direction of the gas distribution pipe is 20 mm. The diameter of the gas distribution pipe is 0.25 m, and the distance between the gas distribution pipe and the bottom of the evaporator is 0.3 m. The heating unit is set 0.4 m above the gas distribution pipe. The heating unit includes two groups of heat exchange tubes, and the heating medium is water vapor.

[0167] The structure of the first gas-liquid separation tank is as follows Figure 5As shown, the first gas-liquid separation tank is provided with three-level baffles, the thickness of the baffle is 4 cm, the length of the baffle is 0.2 m, the bending angle of the baffle is 30°, the distance between the baffle and the liquid phase discharge port N8 at the lower end of the first gas-liquid separation tank is 1.2 m, the surface structure of the baffle is serrated, the tooth height of the sawtooth is 1.5 cm, the tooth pitch of the sawtooth is 1.5 cm, the height difference between the on-site liquid level gauge L1 and the on-site liquid level gauge L2 is 1 m, the on-site pressure gauge P is set 0.8 m above the first gas-liquid separation tank inlet N7, the on-site thermometer T is set 1.3 m above the on-site liquid level gauge L1, and the hand hole H is set 1.4 m below the first gas-liquid separation tank inlet N7.

[0168] According to the present method, the molar ratio of pyridine (purity 99.9%) to hydrogen is 1:10, the hydrogen is pressurized to 3.2 MPa, the pyridine is preheated to 100°C, the preheated pyridine and hydrogen enter the evaporator at a temperature of 140°C and a pressure of 3 MPa, the gas-liquid mixture enters the first pyridine gas-liquid separation tank for separation, and the gaseous phase enters the electric heater at a temperature of 150°C and a pressure of 3 MPa to completely vaporize the pyridine. The vaporized pyridine and hydrogen flow through two tubular fixed-bed reactors equipped with a catalyst and alumina porcelain balls in a ratio of 1:3 (based on the total amount of the catalyst, the metal element palladium content in the catalyst is 1.0wt%, the ruthenium content is 0.8wt%, the carrier alumina content is 98.2wt%, the shape is spherical, and the average particle size of the catalyst is 4 mm). The reaction temperature is controlled at 150°C, the reaction pressure is 3 MPa, and the pyridine mass space velocity is 0.6h -1 The hydrogenation product is cooled to 5°C and separated by the product separation and purification unit. In the second gas-liquid separation tank of the product separation and purification unit, the operating temperature is 5°C and the operating pressure is 2.9MPa. In the flash separation tank, the flash temperature is 5°C and the flash pressure is 0.3MPa. In the distillation tower of the product separation and purification unit, the number of plates is 20 and the distillation pressure is 0.25MPa. The liquid product at the top of the tower is collected for analysis. The conversion rate of pyridine is 99.99%, the piperidine selectivity is 99.9%, the piperidine purity is 99.99%, and the reactor can operate continuously for more than 1200h.

[0169] Example 2

[0170] The production system used is Figure 1 As shown, the evaporation device structure is as follows Figure 2 As shown, the axial length to diameter ratio of the evaporator and the gas distribution pipe is 1:0.7, the gas distribution pipe comprises three groups of through holes, with 45 evenly distributed circular holes, each group of through holes comprises 15 through holes arranged along the axial direction, and each group of through holes is independent of each other. Figure 4Similar to the structure shown, two sets of through holes are symmetrically arranged on the tube wall above the horizontal plane of the gas distribution tube axis, while another set of through holes is arranged below this horizontal plane and staggered with the two sets of through holes above. Specifically, when viewed from above, the through holes below the horizontal plane are located between two adjacent through holes on the horizontal plane. The through hole diameter is 10 mm, and the spacing between any two adjacent through holes along the axial direction of the gas distribution tube is 15 mm. The gas distribution tube has a diameter of 0.25 m, and the distance between the gas distribution tube and the bottom of the evaporator is 0.3 m. The heating unit is located 0.4 m above the gas distribution tube. The heating unit comprises two sets of heat exchange tubes, and the heating medium is water vapor.

[0171] The structure of the first gas-liquid separation tank is as follows Figure 5 As shown, it is the same as in Example 1.

[0172] According to the present method, the molar ratio of pyridine (purity 99.9%) to hydrogen is 1:16, the hydrogen is pressurized to 2.5 MPa, the pyridine is preheated to 100° C., the preheated pyridine and hydrogen enter an evaporator at a temperature of 130° C. and a pressure of 2 MPa; the gas-liquid mixture enters a first gas-liquid separation tank of pyridine for separation, and the gaseous phase enters an electric heater at a temperature of 140° C. and a pressure of 2 MPa to completely vaporize the pyridine. The vaporized pyridine and hydrogen flow through two tubular fixed-bed reactors containing a catalyst and alumina porcelain balls in a ratio of 1:4 (based on the total amount of the catalyst, the metal element palladium content of the catalyst is 0.6 wt%, the nickel content is 0.6 wt%, the carrier silicon oxide content is 98.8 wt%, the shape is spherical, and the average particle size of the catalyst is 3.5 mm). The reaction temperature is controlled at 143° C., the reaction pressure is 2 MPa, and the pyridine mass space velocity is 1.5 h -1 The hydrogenation product is cooled to 3°C and separated by the product separation and purification unit. In the second gas-liquid separation tank of the product separation and purification unit, the operating temperature is 3°C and the operating pressure is 1.9MPa. In the flash separation tank, the flash temperature is 3°C and the flash pressure is 0.3MPa. In the distillation tower of the product separation and purification unit, the number of plates is 20 and the distillation pressure is 0.25MPa. The liquid product at the top of the tower is collected for analysis. The conversion rate of pyridine is 99.99%, the piperidine selectivity is 99.85%, the piperidine purity is 99.9%, and the reactor can operate continuously for more than 1200 hours.

[0173] Example 3

[0174] The parameters of the evaporation device are the same as those in Example 1.

[0175] According to the present method, the molar ratio of pyridine (purity 99.9%) to hydrogen was 1:20, the hydrogen was pressurized to 3.6 MPa, the pyridine was preheated to 100° C., and the preheated pyridine and hydrogen entered the evaporator at a temperature of 140° C. and a pressure of 3.5 MPa; the gas-liquid mixture entered the first pyridine gas-liquid separation tank for separation, and the gaseous phase flow entered the electric heater at a temperature of 155° C. and a pressure of 3.5 MPa to completely vaporize the pyridine. The vaporized pyridine and hydrogen flowed through two tubular fixed-bed reactors containing a catalyst and alumina ceramic balls in a ratio of 1:5 (the catalyst composition and average particle size were the same as in Example 1), and the reaction temperature was controlled at 165° C., the reaction pressure was 3.5 MPa, and the pyridine mass space velocity was 0.3 h -1 The hydrogenation product is cooled to 0°C and separated by the product separation and purification unit. In the second gas-liquid separation tank of the product separation and purification unit, the operating temperature is 0°C and the operating pressure is 2MPa. In the flash separation tank, the flash temperature is 0°C and the flash pressure is 0.4MPa. In the distillation tower of the product separation and purification unit, the number of plates is 20 and the distillation pressure is 0.3MPa. The liquid product at the top of the tower is collected for analysis. The conversion rate of pyridine is 99.99%, the piperidine selectivity is 99.8%, the piperidine purity is 99.9%, and the reactor can operate continuously for more than 1200 hours.

[0176] Example 4

[0177] The parameters of the evaporation device are the same as those in Example 1.

[0178] According to the present method, the molar ratio of pyridine (purity 99.9%) to hydrogen was 1:9, the hydrogen was pressurized to 2.2 MPa, the pyridine was preheated to 100°C, and the preheated pyridine and hydrogen entered the evaporator at a temperature of 110°C and a pressure of 2 MPa; the gas-liquid mixture entered the first pyridine gas-liquid separation tank for separation, and the gaseous phase flow entered the electric heater at 160°C and a pressure of 1.5 MPa to completely vaporize the pyridine. The vaporized pyridine and hydrogen flowed through two tubular fixed-bed reactors equipped with a catalyst and alumina ceramic balls in a ratio of 1:6 (the catalyst composition and average particle size were the same as in Example 1), and the reaction temperature was controlled at 170°C, the reaction pressure was 1.5 MPa, and the pyridine mass space velocity was 2 h -1 The hydrogenation product is cooled to 3°C and separated by the product separation and purification unit. In the second gas-liquid separation tank of the product separation and purification unit, the operating temperature is 3°C and the operating pressure is 1.2MPa. In the flash separation tank, the flash temperature is 3°C and the flash pressure is 0.3MPa. In the distillation tower of the product separation and purification unit, the number of plates is 20 and the distillation pressure is 0.2MPa. The liquid product at the top of the tower is collected for analysis. The conversion rate of pyridine is 99.9%, the piperidine selectivity is 99.2%, the piperidine purity is 99.5%, and the reactor can operate continuously for more than 900 hours.

[0179] Example 5

[0180] The method and evaporation apparatus parameters of Example 1 were followed, except that the gas distribution tube had 21 through-holes, each group of which included seven axially arranged through-holes with a diameter of 20 mm. The spacing between any two adjacent axially aligned through-holes in the gas distribution tube was 30 mm. The pyridine conversion rate was 99.9%, the piperidine selectivity was 99.0%, the piperidine purity was 99.3%, and the reactor was capable of continuous operation for over 800 hours.

[0181] Example 6

[0182] The method and parameters of the evaporation apparatus were followed as in Example 1, except that the molar ratio of pyridine (purity 99.9%) to hydrogen was 1:8, the conversion of pyridine was 99.9%, the piperidine selectivity was 99.4%, the piperidine purity was 99.7%, and the reactor could operate continuously for more than 1000 h.

[0183] Comparative Example 1

[0184] The method of Example 1 was followed, except that the preheated pyridine and hydrogen were directly mixed in a mixer, and the mixed pyridine and hydrogen were heated to 160° C. in an electric heater, followed by a hydrogenation reaction. The pyridine conversion rate was 98%, the piperidine selectivity was 96%, and the piperidine purity was 97.5%. The pyridine conversion rate decreased after the catalyst was operated for 100 h.

[0185] Comparative Example 2

[0186] According to the method and parameters of the evaporation apparatus of Example 1, except that no gas distribution pipe is provided in the evaporator, the conversion rate of pyridine is 99%, the selectivity of piperidine is 97.8%, the purity of piperidine is 99.0%, and the reactor can operate continuously for about 500 hours.

[0187] The above results show that, compared with the comparative example, the embodiment of the piperidine production system and method of the present invention adopts continuous production, mild conditions, good catalyst stability, long service life, green production process, less three wastes generated, and hydrogen can be recycled with low consumption. After the pyridine is completely vaporized, it enters the hydrogenation reactor and can fully contact with the catalyst, thereby improving the pyridine conversion rate, piperidine selectivity and purity, reducing the degree of catalyst deactivation, and increasing the operating time of the reactor, thereby solving the problems of low pyridine conversion rate, piperidine selectivity and piperidine purity in the piperidine production process.

[0188] The preferred embodiments of the present invention have been described in detail above, but the present invention is not limited thereto. Within the technical concept of the present invention, various simple variations of the technical solution of the present invention may be made, including combining the various technical features in any other appropriate manner. These simple variations and combinations should also be regarded as disclosed in the present invention and fall within the scope of protection of the present invention.

Claims

1. A piperidine production system, characterized in that, The production system comprises a feeding unit, a vaporizing unit, a reaction unit, a cooling unit and a product separation and purification unit connected in sequence; Wherein, the feeding unit comprises a circulating hydrogen buffer tank, a hydrogen compressor and a hydrogen buffer tank connected in sequence; Wherein, the vaporization unit comprises an evaporator, a first gas-liquid separation tank and an electric heater connected in sequence; From bottom to top along the axial direction of the evaporator, the evaporator includes a distribution unit and a heating unit in sequence. The distribution unit includes a gas distribution pipe arranged non-parallel to the axial direction of the evaporator. The pipe wall of the gas distribution pipe is provided with a through hole. One end of the gas distribution pipe is provided with a gas phase inlet. The bottom of the evaporator is provided with a liquid phase inlet, so that the gas phase passes through the gas distribution pipe, mixes with the liquid phase, and enters the heating unit; The first gas-liquid separation tank is provided with a baffle; The electric heater is connected to the gas phase outlet of the first gas-liquid separation tank and is used to heat the gas phase flow from the first gas-liquid separation tank into superheated steam; Wherein, the reaction unit includes a hydrogenation reactor for hydrogenating hydrogen and pyridine; Wherein, the cooling unit comprises a heat exchanger and a cooler connected in sequence, which are optionally used to cool the liquid phase flow from the reaction unit; The product separation and purification unit is used to separate and purify the liquid phase flow from the cooling unit to obtain piperidine.

2. The production system according to claim 1, wherein: The production system further comprises a preheating unit connected to the raw material vaporization unit; Preferably, the preheating unit comprises a pyridine preheater; And / or, the gas phase outlet of the second gas-liquid separation tank is connected to the circulating hydrogen buffer tank; And / or, the reaction unit comprises two hydrogenation reactors, and the hydrogenation reactors are shell-and-tube fixed-bed reactors; Preferably, the two hydrogenation reactors are connected in series; And / or, the product separation and purification unit includes a second gas-liquid separation tank, a flash separation tank and a distillation tower connected in sequence.

3. The production system according to claim 1 or 2, wherein: The distribution unit of the evaporator includes a gas distribution pipe and a support beam for fixing the gas distribution pipe, and the gas distribution pipe is perpendicular to the axial direction of the evaporator; Preferably, along the axial direction of the gas distribution pipe, there is a gap between the gas distribution pipe and the right side wall of the evaporator for liquid phase circulation; Preferably, the gas phase stream is introduced from the gas phase inlet (N4); Preferably, the diameter of the gas distribution pipe is 0.1-0.8m; Preferably, the distance between the gas distribution pipe and the bottom of the evaporator is 0.05-0.6m; Preferably, n groups of through holes are provided along the circumference of the wall of the gas distribution pipe, each group independently including m through holes arranged along the axial direction, n≥3, m≥3; two adjacent groups of through holes are preferably staggered; Preferably, no through-hole is provided on the pipe wall above the horizontal plane where the axis of the gas distribution pipe is located; Preferably, the through hole has a diameter of 6-30 mm, preferably 8-18 mm; Preferably, the number of the through holes is 9-50, preferably 20-45; Preferably, the distance between any two adjacent through holes along the axial direction of the gas distribution pipe is 10-36 mm, preferably 15-30 mm; and / or, the axial length to diameter ratio of the evaporator and the gas distribution pipe is 1:0.5-0.9; Preferably, the shape of the through hole is selected from at least one of a circular hole, a triangular hole, a square hole, a polygonal hole, an elliptical hole, and a diamond hole, and is preferably a circular hole.

4. The production system according to any one of claims 1 to 3, wherein: The heating unit is arranged 0.1-1.2m above the gas distribution pipe; Preferably, the heating unit comprises at least one set of heat exchange tubes for circulating the heating medium; Preferably, the heating medium of the heating unit is water vapor; Preferably, the evaporator is provided with a heating medium inlet (N2) at the top and a heating medium outlet (N3) at the bottom, so that the heating medium enters the evaporator from the heating medium inlet (N2) and flows out from the heating medium outlet (N3).

5. The production system according to any one of claims 1 to 4, wherein: The first gas-liquid separation tank is provided with 1-5 levels of baffles, preferably 2-4 levels of baffles; Preferably, the thickness of the baffle is 2-10 cm, preferably 3-8 cm; Preferably, the length of the baffle is 0.1-0.6 m, preferably 0.2-0.5 m; Preferably, the baffle has a bending angle of 10-50°, preferably 10-30°; Preferably, the distance between the baffle and the liquid phase discharge port (N8) at the lower end of the first gas-liquid separation tank is 0.6-2m, preferably 0.8-1.4m; Preferably, the surface structure of the baffle is serrated; Preferably, the saw teeth have a tooth height of 1-5 cm, preferably 1-3 cm; Preferably, the tooth pitch of the saw teeth is 1-5 cm, preferably 1-3 cm; And / or, the first gas-liquid separation tank is provided with an on-site liquid level meter (L1) and an on-site liquid level meter (L2); Preferably, the height difference between the on-site liquid level gauge (L1) and the on-site liquid level gauge (L2) is 0.4-1.4 m, preferably 0.6-1.2 m; and / or, the first gas-liquid separation tank is provided with an on-site pressure gauge (P); Preferably, the on-site pressure gauge (P) is arranged 0.3-1.8 m above the first gas-liquid separation tank inlet (N7), preferably 0.5-1.4 m; And / or, the first gas-liquid separation tank is provided with an on-site thermometer (T); Preferably, the on-site thermometer (T) is arranged 0.5-2m above the on-site liquid level meter (L1), preferably 0.8-1.6m; and / or, the first gas-liquid separation tank is provided with a hand hole (H); Preferably, the hand hole (H) is arranged 0.8-2.5m below the first gas-liquid separation tank inlet (N7), preferably 1-2m.

6. A method for continuously producing piperidine using the production system according to any one of claims 1 to 5, characterized in that: The method includes: (1) The hydrogen is passed into the circulating hydrogen buffer tank, pressurized by the hydrogen compressor, and then enters the hydrogen buffer tank; (2) introducing hydrogen and pyridine in the hydrogen buffer tank into a raw material vaporization device for vaporization to obtain a vaporized raw material; The vaporization comprises introducing hydrogen from the gas phase inlet (N4) and pyridine from the liquid phase inlet (N5) at the bottom of the evaporator, so that the hydrogen passes through the gas distribution pipe, mixes with the pyridine, and enters the heating unit for a first heating; sending the heated logistics into a first gas-liquid separation tank for gas-liquid separation; and sending the gas phase logistics obtained by the gas-liquid separation into an electric heater for a second heating; (3) passing the vaporized feedstock into a hydrogenation reactor, and performing a hydrogenation reaction in the presence of a catalyst and optionally a diluent to obtain a hydrogenated product; (4) Cooling and separating the hydrogenation product.

7. The method according to claim 6, wherein: The molar ratio of hydrogen to pyridine is 8-40:1, preferably 10-30:1; And / or, the mass ratio of the catalyst to the diluent is 1:1-10, preferably 1:2-5; Preferably, the catalyst comprises a metal element and a carrier; Preferably, based on the total amount of the catalyst, the content of the metal element is 1-2 wt%, and the content of the carrier is 98-99 wt%; Preferably, the metal element is selected from at least one of platinum, rhodium, palladium, nickel, copper, zinc, ruthenium and molybdenum, preferably at least one of palladium, nickel and ruthenium; Preferably, the carrier is selected from at least one of alumina, silica, activated carbon and molecular sieve, preferably alumina; Preferably, the average particle size of the catalyst is 3-5 mm; Preferably, the catalyst is spherical in shape; And / or, the diluent is selected from at least one of quartz sand, alumina ceramic balls, silicon oxide ceramic balls and metal ceramic balls.

8. The method according to claim 6 or 7, wherein: The separation in step (4) is carried out in the second gas-liquid separation tank, the flash separation tank and the distillation tower; Preferably, the hydrogen obtained from the second gas-liquid separation tank is returned to the circulating hydrogen buffer tank; And / or, the step (2) further includes preheating the pyridine in a pyridine preheater.

9. The method according to any one of claims 6 to 8, wherein: The pressurized pressure in step (1) is 1-4 MPa; And / or, the first heating conditions include: a temperature of 100-180° C., preferably 100-150° C.; a pressure of 1-5 MPa, preferably 1.5-4 MPa; And / or, the second heating conditions include: a temperature of 100-200° C., preferably 130-180° C.; a pressure of 1-5 MPa, preferably 1.5-4 MPa; And / or, the hydrogenation conditions in step (3) include: temperature of 120-180°C, pressure of 1-4 MPa, pyridine mass space velocity of 0.1-2h -1 .

10. The method according to any one of claims 6 to 9, wherein: The cooling temperature is 0-20°C.

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