CN108069815B - Method for preparing high-purity butane - Google Patents
Method for preparing high-purity butane Download PDFInfo
- Publication number
- CN108069815B CN108069815B CN201610991238.4A CN201610991238A CN108069815B CN 108069815 B CN108069815 B CN 108069815B CN 201610991238 A CN201610991238 A CN 201610991238A CN 108069815 B CN108069815 B CN 108069815B
- Authority
- CN
- China
- Prior art keywords
- reaction
- temperature
- heat
- liquefied gas
- reactor
- Prior art date
- Legal status (The legal status is an assumption and is not a legal conclusion. Google has not performed a legal analysis and makes no representation as to the accuracy of the status listed.)
- Active
Links
Images
Classifications
-
- C—CHEMISTRY; METALLURGY
- C07—ORGANIC CHEMISTRY
- C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
- C07C5/00—Preparation of hydrocarbons from hydrocarbons containing the same number of carbon atoms
- C07C5/02—Preparation of hydrocarbons from hydrocarbons containing the same number of carbon atoms by hydrogenation
- C07C5/03—Preparation of hydrocarbons from hydrocarbons containing the same number of carbon atoms by hydrogenation of non-aromatic carbon-to-carbon double bonds
-
- C—CHEMISTRY; METALLURGY
- C07—ORGANIC CHEMISTRY
- C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
- C07C2523/00—Catalysts comprising metals or metal oxides or hydroxides, not provided for in group C07C2521/00
- C07C2523/70—Catalysts comprising metals or metal oxides or hydroxides, not provided for in group C07C2521/00 of the iron group metals or copper
- C07C2523/76—Catalysts comprising metals or metal oxides or hydroxides, not provided for in group C07C2521/00 of the iron group metals or copper combined with metals, oxides or hydroxides provided for in groups C07C2523/02 - C07C2523/36
- C07C2523/84—Catalysts comprising metals or metal oxides or hydroxides, not provided for in group C07C2521/00 of the iron group metals or copper combined with metals, oxides or hydroxides provided for in groups C07C2523/02 - C07C2523/36 with arsenic, antimony, bismuth, vanadium, niobium, tantalum, polonium, chromium, molybdenum, tungsten, manganese, technetium or rhenium
- C07C2523/85—Chromium, molybdenum or tungsten
- C07C2523/888—Tungsten
Abstract
The invention discloses a method for preparing high-purity butane. The method comprises the following steps: providing a shell and tube reactor: the tube pass of the reactor is filled with a hydrogenation catalyst, and the shell pass of the reactor comprises a heat-conducting carrier decane for circularly taking heat; controlling the temperature of the circulating decane in the shell pass to be 205 +/-20 ℃ and the pressure to be 0.13 MPa-0.21 MPa; preheating raw material liquefied gas and hydrogen to a feeding temperature, entering a tube pass of a reactor, and carrying out contact reaction with a hydrogenation catalyst; and separating and fractionating the reaction effluent to obtain a butane product. The invention utilizes the characteristic that the heat of vaporization of the heat-conducting medium is far more than the sensible heat at the same temperature, can quickly absorb the reaction heat, and thus, can effectively, quantitatively and accurately control the temperature of the reaction hot point. By utilizing the characteristic of large reaction equilibrium constant of the reaction at low temperature, the invention can produce high-purity butane, has stable product quality and ensures the safe and stable operation of the hydrogenation process of liquefied gas.
Description
Technical Field
The invention relates to a butene hydrogenation production technology, in particular to a method for producing high-purity butane by hydrogenation of butene.
Background
Isobutane is a basic chemical raw material, and dehydrogenation can produce isobutene; it can produce Propylene Oxide (PO) and coproduce Tertiary Butyl Alcohol (TBA) and the like by carrying out a co-oxidation reaction with propylene.
A mixed C4 which is a byproduct of an ethylene unit of a chemical plant, a coking unit of an oil refinery and a catalytic cracking unit and contains isobutane. The isobutane obtained from the mixed C4 separation contains more or less a certain amount of olefins and therefore needs to be further hydrofinished before it can be used as a chemical feedstock.
CN1160701A introduces C3Method for hydrogenating distillate, but the method is aimed at making C3The acetylenes in the distillate are selectively hydrogenated, not the entire distillate (including the mono-olefins). CN1145891A describes a hydrogenation process, but this process is only applicable to C5And (4) hydrogenating the distillate to prepare pentane. When a non-noble metal hydrogenation catalyst is used, the conversion rate of mono-olefin is low. CN01114163.8 describes a method for preparing liquefied petroleum gas for vehicles by hydrogenation of liquefied petroleum gas, but the method has the disadvantage that the index of the product after hydrogenation is olefin<5.0%。
The CN102311760A and CN102311783A patents disclose a method for preparing ethylene cracking material by hydrogenation of liquefied petroleum gas. The method is applicable to C4Fraction hydrogenation process, but the product index after hydrogenation is lower, and the olefin content in the product after hydrogenation is not more than 1.0 percent.
Isobutane and propylene are subjected to a co-oxidation reaction to produce Propylene Oxide (PO) and co-produce Tertiary Butyl Alcohol (TBA), one of raw materials used in the technology is iso-butane, and the olefin content in the iso-butane is required to be less than 50mg-3The technical index has high requirement. The olefin content in the obtained product reaches 0.1 percent (namely 1000 mg.m) by adopting a conventional hydrogenation method-3) It is quite difficult. It is difficult for the conventional hydrogenation method to satisfy the requirement for producing high-purity butane.
C4The olefin hydrogenation reaction is an equilibrium reaction, and when the reaction temperature is low, the reaction speed is too low, and the technology has no practical value; at high reaction temperatureThe reaction rate increases, but the equilibrium constant is too small. Due to the limitation of reaction equilibrium at high temperature, the alkene content in the hydrogenated product is high, and high-purity C cannot be obtained4An alkane; in addition, the reaction is a strongly exothermic reaction, which results in a narrow temperature range for effective use of the reaction. It follows that how to control the reaction temperature within an effective range of use is a difficulty in achieving this technology. Also for the reasons mentioned above, there has been no mature process for producing butane with high purity by hydrogenation so far.
In the prior art, in order to improve the hydrogenation effect of the liquefied petroleum gas, the hydrogenation reaction has to be carried out at higher temperature. Since the olefin hydrogenation reaction has a severe heat release, the reaction overtemperature is easily caused by increasing the reaction temperature, and once the control is not good, irreversible damage on the performance of the catalyst is brought. Thus, the prior art does not provide a good solution for producing (iso) butane at high purity.
Disclosure of Invention
Aiming at the defects of the prior art, the invention innovates in the aspect of production process and provides a method for producing high-purity butane. The method can control the reaction temperature stably without over temperature, and the production device runs stably and the product quality is stable.
The inventors of the present application, by making C4The hydrogenation reaction of the fraction (liquefied gas) is subjected to system experiments and deep research, and the following experimental research results and knowledge are obtained:
(1) the raw material liquefied gas has the characteristics that: one is that the sulfur content is high, and the content is 0.02%; the second is 5.212% of alkene content.
TABLE 1 liquefied gas composition
| Serial number | Components | Composition in wt% |
| 1 | Methane | 0.113 |
| 2 | Ethane (III) | 0.356 |
| 3 | Propane | 9.674 |
| 4 | 1-butene | 3.652 |
| 5 | Isobutene | 1.267 |
| 6 | Cis-2-butene | 0.042 |
| 7 | Trans-2-butene | 0.251 |
| 8 | N-butane | 24.523 |
| 9 | Isobutane | 58.998 |
| 10 | Isopentane | 1.104 |
| 11 | Sulfur content | 0.02 |
| Total up to | 100 | |
| Olefin subtotal | 5.212 |
(2) The reaction characteristics are as follows: the olefin hydrogenation reaction is a strongly exothermic reaction; the liquefied gas in Table 1 is used as a raw material to carry out hydrogenation reaction, and the reaction heat when all the olefins are hydrogenated is 26.13 KJ/mol; the adiabatic reaction temperature rise can reach 29 ℃.
(3) The result of the kinetic study was: the reaction is a fast reaction and thus the exotherm in the reaction bed is non-uniform. When the reaction is carried out through the catalyst bed layer, the reaction heat evolution reaches about 80 percent of the whole reaction heat when the reaction residence time is 1/3 of the whole residence time. Therefore, a "temperature bag" or a hot spot exists at the position of the inlet section of the reaction bed layer, which is far from the inlet 1/3, and the temperature bag can be reduced only by selecting the conventional multi-section adiabatic reactor or the tubular reactor, and the effect is not obvious.
In addition, the reaction kinetics research result also shows that the excessive hydrogen amount (the excessive hydrogen/liquefied gas volume ratio) can increase the gasification rate of the liquefied gas and the linear velocity of the reaction feeding material, so the excessive hydrogen amount is unfavorable for the reaction.
(4) The result of the thermodynamic study is as follows: as can be seen from table 1, isobutene, which is the highest olefin of C4 in the liquefied gas, was calculated for the reaction equilibrium constant, and the calculation results are shown in table 2.
TABLE 2 reaction temperature vs. equilibrium constant relationship
| Serial number | Reaction temperature/. degree.C | Reaction equilibrium constant/ |
| 1 | 160 | 1.96E+10 |
| 2 | 170 | 9.43E+09 |
| 3 | 180 | 4.67E+09 |
| 4 | 190 | 2.38E+09 |
| 5 | 200 | 1.25E+09 |
| 6 | 210 | 6.73E+08 |
| 7 | 220 | 3.71E+08 |
| 8 | 230 | 2.10E+08 |
| 9 | 240 | 1.21E+08 |
| 10 | 250 | 7.14E+07 |
| 11 | 260 | 4.29E+07 |
| 12 | 270 | 2.62E+07 |
| 13 | 280 | 1.63E+07 |
| 14 | 290 | 1.03E+07 |
| 15 | 300 | 6.65E+06 |
| 16 | 310 | 4.34E+06 |
| 17 | 320 | 2.88E+06 |
| 18 | 330 | 1.93E+06 |
| 19 | 340 | 1.31E+06 |
As can be seen from the data in Table 2, the equilibrium constant rapidly decreases with increasing reaction temperature. For example, the reaction temperature is increased from 210 ℃ to 220 ℃, the temperature is increased by only 10 ℃, and the equilibrium constant is reduced by nearly 50%.
Within the calculated temperature range, the Δ G θ m of the olefin hydrogenation reaction is less than 0, namely, the olefin hydrogenation reaction can be spontaneously carried out under the corresponding temperature condition; the K θ eq values for each reaction were very large, indicating that the reaction proceeded easily and decreased with increasing temperature, indicating that increasing the reaction temperature was detrimental to the performance of each reaction.
The temperature is an important factor influencing the low-carbon olefin hydrogenation reaction, and when the subsequent olefin hydrogenation reaction condition is selected, the reaction pressure is increased, and the reaction temperature is reduced to reduce the influence of the temperature on the reaction balance; secondly, the reaction temperature is controlled to be within an effective use range as much as possible, so that the olefin hydrogenation reaction is carried out to a certain depth.
From the thermodynamic viewpoint, a low temperature is favorable for this reaction, but if the reaction temperature is low, the reaction speed is slow although the reaction equilibrium constant is large; on the contrary, if the reaction temperature is high, the reaction equilibrium constant is small although the reaction rate is high. WhileThe reaction requires high purity of hydrogenated product (product index is less than 50 mg.m)-3) Therefore, if the diffusion of the reaction heat is not well controlled (or the reaction temperature rise is not effectively controlled) for this reaction, the reaction control is shifted to thermodynamic control, and the conversion rate of the reaction is reduced.
(5) The investigation result of the reaction process conditions is as follows: in the reaction, an FH-10B catalyst is selected, the catalyst is a high-activity low-carbon olefin hydrofining catalyst which is newly developed and produced by the research institute of petrochemical engineering, and the physical properties of the catalyst are listed in Table 3.
TABLE 3 physical and chemical Properties of LH-10B catalyst
| Detecting items | Quality index | Analytical method |
| Reactive metal | W-Mo-Co | Colorimetric method |
| Metal content, wt.% | 15～23 | |
| Pore volume/(mL. g)-1) | ≮0.42 | Low temperature nitrogen adsorption |
| Specific surface area/(m)2·g-1) | ≮220 | Low temperature nitrogen adsorption |
| Side pressure strength/(N cm)-1) | ≮150 | Progressive intensity meter |
| Outer shape | Clover type | Visual inspection of |
| Bulk density/(g. mL)-1) | 0.75～0.85 | Measuring cylinder method |
After the catalyst was determined, we investigated the reaction conditions and combined the results of the reaction conditions investigation experiment and thermal study to obtain the reaction conditions and results suitable for the reaction shown in table 4.
TABLE 4 reaction conditions and results
| Reaction conditions | |
| Reaction pressure/MPa | 8.0 |
| Liquefied gas feeding volume airspeed/h-1 | 0.2～1.5 |
| Hydrogen/liquefied gas, v/v | 200～300 |
| Inlet temperature/. degree.C | 180～190 |
| Maximum temperature point of bed (hot spot)/° c | ≯210 |
| Average reaction temperature/. degree.C | 140～340 |
| Results of the reaction | |
| Olefin content in the reaction product, mg.m-3 | ≯50 |
The results of the above studies suggest to us that:
(1) suitable reaction conditions should be high pressure and low temperature. The influence of the temperature on the reaction balance is reduced as much as possible by increasing the reaction pressure and reducing the reaction temperature.
(2) Since the reaction requires high purity of the hydrogenated product (the product index is less than 50 ppm), the reaction temperature control is the main control condition after the pressure selection, namely the problem of taking out the reaction heat is the main problem to be solved.
(3) The reaction equilibrium constant can be improved and the conversion rate can be improved by reducing the 'temperature bulb' of the reaction bed layer.
Based on the above findings, the inventors of the present application propose a method for preparing high-purity butane by hydrogenating liquefied gas, which comprises the following steps:
(1) providing a shell and tube reactor: the tube pass of the reactor is filled with a hydrogenation catalyst, and the shell pass of the reactor comprises a heat-conducting carrier decane for circularly taking heat; controlling the temperature of the circulating decane in the shell pass to be 205 +/-20 ℃ and the pressure to be 0.13 MPa-0.21 MPa;
(2) preheating raw material liquefied gas and hydrogen to a feeding temperature, entering a tube pass of a reactor, and carrying out contact reaction with a hydrogenation catalyst; and separating and fractionating the reaction effluent to obtain a butane product.
In step (1) of the present invention, the tubular reactor is a conventional reactor in the art. Which typically includes a tube side and a shell side. The pressure of the circulating decane is controlled to be 0.13MPa to 0.21 MPa. Wherein, the purity of the decane as the heat-conducting medium is generally required to be more than 90wt percent, and preferably more than 95wt percent.
In the invention, when the pressure is set to be 0.207MP, decane is in a liquid phase when the temperature is 205 ℃ lower than the gasification temperature; decane is in the vapour phase at a temperature 205 ℃ above the vaporisation temperature. Under the set pressure range and circulation temperature, the heat conducting carrier decane selected by the invention is liquid under the condition that the catalyst bed layer in the reaction tube is at normal reaction temperature (the reaction hot spot is less than 195 ℃), the reaction heat released in the reaction tube is taken out through the heat carrier circulated by the shell layer (pass), and the sensible heat of the circulating heat carrier is utilized at the moment (the sensible heat of decane at 195 ℃ is 0.403KJ/mol, and the sensible heat of decane at 205 ℃ is 0.411 KJ/mol). When the catalyst bed layer in the reaction tube is at an over-temperature (a certain section or local hot spot of the catalyst bed layer is over-temperature, and is more than or equal to 205 ℃), the temperature at the moment just reaches the vaporization temperature of decane in the shell layer of the reactor, so that the decane in the shell layer is rapidly vaporized, liquid is changed into vapor, and the heat taking process utilizes the latent heat of decane as a heat carrier.
The invention provides the technical scheme through the deep research on the technical process of butane production by hydrogenation of liquefied gas. The method of the invention has the following characteristics: (1) the heat extraction process utilizes the principle that the latent heat of the same substance is greater than the sensible heat to quickly absorb the reaction heat, so that the hot spot of the reaction bed layer is eliminated as soon as possible. Since the sensible heat of decane at 205 ℃ is 0.411KJ/mol and the heat of vaporization is 39.397 KJ/mol, the heat of vaporization at this temperature is 88.55 times that of the sensible heat. (2) The effective quantitative control of the temperature of the hot spot of the reaction is achieved by utilizing the characteristic that the vaporization temperature of a substance is certain under certain pressure and can be vaporized once reaching the vaporization temperature. (3) The vaporized gas is recycled after condensation and cooling.
The preheating temperature of the hydrogen and the liquefied gas in the method is generally 140-240 ℃, and preferably 180-190 ℃; the highest temperature point (i.e. hot spot) of the catalyst bed in the tube pass can reach 190-240 ℃, and preferably 200-210 ℃; the reaction pressure is generally 4MPa to 10MPa, preferably 7MPa to 9 MPa; the volume ratio of the hydrogen gas to the liquefied gas is 200-1000, preferably 200-300; the volume airspeed of the liquefied gas feeding (liquid) is 0.3-2.0 h-1Preferably 0.6 to 0.8 h-1。
The hydrogenation catalyst used in the process of the present invention may be selected from conventional hydrogenation catalysts in the art. The hydrogenation catalyst can be a supported hydrogenation catalyst or a bulk phase hydrogenation catalyst. The supported hydrogenation catalyst comprises a carrier and a supported active metal component. The support is typically a porous refractory inorganic oxide or activated carbon. In particular, the support is generally selected from Al2O3SiO-containing2Al of (2)2O3、TiO2Molecular sieve-containing Al2O3And activated carbon. The active metal component is selected from a noble metal or a non-noble metal. The noble metal typically comprises one or more of Pt, Pd and Re, and the non-noble metal typically is selected from one or more of W, Mo, Ni and Co. In the present invention, the non-noble metal preferably comprises W and/or Ni. The content of the noble metal is generally 0.1-2.0 wt% calculated by metal oxide; the content of non-noble metal components is generally from 5% to 35% by weight.
The bulk phase hydrogenation catalyst contains three metal components of Mo, W and Ni, wherein W, Ni exists in a form of composite oxide: nixWyOzZ ═ x +3y, Mo exists in the oxide form: MoO3. Composite oxide NixWyOzThe ratio (atomic mol ratio) of the x and the y is 1: 8-8: 1, preferably 1: 4-4: 1. Compound medicineDouble oxide NixWyOzAnd oxide MoO3The weight ratio of (A) to (B) is 1: 10-10: 1, preferably 1: 5-5: 1. Composite oxide Ni in bulk phase catalystxWyOzAnd oxide MoO3The total weight content of (A) is 40-100%, preferably 50-80%. (the above catalyst composition is in the oxidized state and the catalyst is used by sulfiding according to methods well known to those skilled in the art).
Compared with the prior art, the method has the following outstanding technical effects:
1. the adoption of the vulcanization type catalyst solves the problem of poisoning effect of trace sulfur in the raw material on the traditional noble metal catalyst in the hydrogenation reaction process, reduces the use cost of the catalyst and prolongs the operation period of the catalyst.
2. By utilizing the characteristic that the vaporization (phase change) heat of the heat-conducting medium decane is far greater than the sensible heat at the same temperature, the reaction heat can be quickly absorbed, the hot point of a reaction bed layer is eliminated as soon as possible, and the safe and stable operation of the liquefied gas hydrogenation process is ensured.
3. By utilizing the characteristic that the vaporization temperature of the heat-conducting medium decane is certain under certain pressure, once the vaporization temperature of the heat-conducting medium decane is reached, the decane can be rapidly vaporized, so that the temperature of a reaction hot spot can be effectively, quantitatively and accurately controlled.
4. The vaporization heat of decane as heat conducting medium is utilized fully to control the temperature of the reaction bed layer and avoid over temperature, so that the hydrogenation reaction is operated stably at relatively low temperature.
Drawings
FIG. 1 is a schematic diagram of a process flow for producing high-purity butane by hydrogenation of liquefied gas.
Wherein: 1-hydrogen, 2-hydrogen feeding heater, 3-liquefied gas, 4-liquefied gas feeding heater, 5-tubular reactor, 6-reactor shell circulating liquid level, 7-reactor tubular catalyst bed, 8-reactor shell gas phase extraction outlet, 9-reactor shell liquid phase extraction outlet, 10-reactor shell gas phase extraction material condenser, 11-condensate storage tank and 12-reactor shell gas phase extraction line; 13-a condensate liquid extraction line, 14-a reactor shell layer circulating liquid storage tank, 15-a reactor shell layer circulating liquid extraction line, 16-a reactor shell layer circulating liquid cooler, 17-a reactor shell layer circulating liquid storage tank extraction line, 18-a circulating liquid pump, 19-a circulating liquid inlet and 20-a reactor extraction line.
Detailed Description
The method of the present invention is described in more detail below with reference to the accompanying drawings.
As shown in fig. 1, the process flow of the method of the present invention comprises:
(1) heating of raw materials
The reaction raw materials of hydrogen 1 and liquefied gas 2 are heated to a certain temperature by a hydrogen feeding heater 3 and a liquefied gas feeding heater 4 respectively and then enter a tubular reactor 5.
(2) Hydrogenation reaction and heat extraction
The temperature of the circulating decane in the shell layer of the reactor is generally set at 190 ℃ and the pressure is set at 0.207MPa, at which the decane is in liquid phase at a temperature lower than 205 ℃ and in vapor phase at a temperature higher than 205 ℃. The selected heat-conducting carrier is liquid under the condition that the catalyst bed layer in the reaction tube is at normal reaction temperature (the hot point is less than 195 ℃), the reaction heat released in the reaction tube is taken out by a shell layer circulating heat carrier, and the sensible heat of the circulating heat carrier is utilized at the moment. When the catalyst bed layer in the reaction tube is at an over-temperature (the over-temperature of a certain section or local hot spot of the catalyst bed layer is more than 205 ℃), the temperature at this moment just reaches the vaporization temperature of decane in the shell layer of the reactor, so that the decane in the shell layer is rapidly vaporized, liquid is changed into vapor, and the heat extraction process utilizes the latent heat of a heat carrier at this moment. The control of the temperature of the reaction bed by the heat of vaporization of decane was chosen because: 1) the latent heat of the same substance is greater than the sensible heat, and the reaction heat can be quickly absorbed, so that the hot spot of the reaction bed layer can be eliminated as soon as possible. 2) The effective quantitative control of the temperature of the hot spot of the reaction is achieved by utilizing the characteristic that the vaporization temperature of a substance is certain under certain pressure and can be vaporized once reaching the vaporization temperature.
The process of the present invention is further described below by means of specific examples.
Comparative examples 1 to 2
The LH-10B catalyst is adopted, the reaction raw materials are the raw materials in the table 1, the reaction process flow is as shown in the figure 1, the shell layer heating medium of the shell layer reactor is heat conduction oil, the technical specification of the heat conduction oil is shown in the table 5, and the reaction conditions and results of the comparative examples 1 and 2 are shown in the table 6.
TABLE 5 great wall L-QB300 thermal oil technical Specification
| Name of item | |
| Initial boiling point/. degree.C | 348 |
| Flash point/. degree.C | 226 |
| Distillation range, deg.C | |
| HK/2%/97% | 348/376/473 |
| Pour point/. degree.C | -12 |
| Etching of | 1 |
| Density, kg/m3 | 868.5 |
| Appearance of the product | Transparent, without suspended matter and precipitate |
| Deterioration rate,% (m/m) is not more than | 5.5 |
Note: 1. the great wall L-QB300 heat conducting oil is prepared by adding various additives such as cleaning, dispersing, high oxidation resistance and the like into refined narrow-fraction mineral base oil; 2. the great wall L-QB300 heat conducting oil is suitable for a forced or non-forced circulation closed heat transfer system with the maximum temperature not more than 300 ℃; 3. the technical specification is that the product meets the following specifications: SH/T0677-1999; 4. the great wall L-QB300 heat conducting oil is suitable for forced or non-forced circulation closed heat transfer system, and may be used in heating, drying and other processes, such as timber processing, textile dyeing and finishing, food processing, chemical industry and other industry.
TABLE 6 reaction conditions and results
| Comparative example 1 | Comparative example 2 | |
| Reaction conditions | ||
| Reaction pressure/MPa | 8.0 | 9.0 |
| Liquefied gas feeding volume airspeed/h-1 | 0.6 | 0.6 |
| Hydrogen/liquefied gas, v/v | 300 | 200 |
| Reaction inlet temperature/. degree.C | 185 | 195 |
| Bed temperature rise/. degree.C | 27 | 28 |
| Maximum temperature point of bed (hot spot)/° c | 212 | 228 |
| Average reaction temperature/. degree.C | 198.5 | 214 |
| Results of the reaction | ||
| Olefin content in the reaction product, mg.m-3 | 200 | 320 |
Note: 1. the shell layer of the reactor is internally provided with L-QB300 heat conducting oil with the pressure of 1.0 MPa; 2. the reaction product was analyzed as a gas phase.
As can be seen from the data in Table 6, the reaction temperature is controlled by using conventional heat transfer oil, and the alkene content in the hydrogenated product is far higher than 50mg.m-3And the technical requirement of high-purity butane cannot be met.
Examples 1 to 3
The LH-10B catalyst is adopted in the reaction, the raw materials are shown in the table 1, the reaction process flow is the same as that of the tubular reactor shown in the figure 1, decane (purity 98 wt%) is used as a shell heat-conducting medium of the tubular reactor, and the reaction conditions and results of the examples 1-3 are shown in the table 7.
Table 7 main operating conditions and results of liquefied gas hydrogenation reactor
| Item | Example 1 | Example 2 | Example 3 |
| Reaction conditions | |||
| Reactor pressure/MPa | 7.0 | 8.0 | 9.0 |
| Molar ratio of hydrogen to oil of reaction | 300 | 200 | 350 |
| Volume space velocity/h of liquefied gas without hydrogenation-1 | 0.6 | 0.7 | 0.8 |
| Reactor inlet temperature/. degree.C | 180 | 185 | 190 |
| Reactor hot spot temperature/. degree.C | 184 | 190 | 194 |
| Temperature rise/deg.C of the reactor | 4 | 5 | 4 |
| Average reaction temperature/. degree.C | 182 | 187.5 | 192 |
| The shell of the reactor is controlled at temperature/° c | 180 | 185 | 190 |
| Reactor shell control pressure/MPa | 0.15 | 0.17 | 0.20 |
| Results of the reaction | |||
| Olefin content in the reaction product, mg.m-3 | 42 | 36 | 34 |
Note: 1. the reaction product was analyzed as a gas phase.
As can be seen from the data in Table 7, when decane is used as the shell heat-conducting medium of the tubular reactor, the temperature of the reaction bed layer is relatively uniform, the temperature rise of the bed layer is relatively small, and the reaction bed layer is not over-temperature, which shows that the reaction temperature control scheme of the invention is reasonable and effective in controlling the reaction temperature, and the alkene content in the hydrogenated product is less than 501.0mg.m-3And the technical requirement of high-purity butane is met.
Claims (11)
1. A method for preparing high-purity butane by hydrogenation of liquefied gas comprises the following steps:
(1) providing a shell and tube reactor: the tube pass of the reactor is filled with a hydrogenation catalyst, and the shell pass of the reactor comprises a heat-conducting carrier decane for circularly taking heat; controlling the temperature of the circulating decane in the shell pass to be 205 +/-20 ℃ and the pressure to be 0.13 MPa-0.21 MPa;
(2) preheating raw material liquefied gas and hydrogen to a feeding temperature, entering a tube pass of a reactor, and carrying out contact reaction with a hydrogenation catalyst; and separating and fractionating the reaction effluent to obtain a butane product.
2. The method of claim 1, wherein the pre-heating temperature of the liquefied gas and hydrogen is 140 ℃ to 240 ℃.
3. The method of claim 2, wherein the pre-heating temperature of the liquefied gas and hydrogen is 180 ℃ to 190 ℃.
4. The method according to claim 1, wherein the reaction conditions in step (2) are as follows: the reaction pressure is 4MPa to 10MPa, the volume ratio of hydrogen to liquefied gas is 200 to 1000, and the feeding volume airspeed of the liquefied gas is 0.3 to 2.0 h-1。
5. The method according to claim 4, wherein the reaction conditions in step (2) are as follows: the reaction pressure is 7MPa to 9MPa, the volume ratio of hydrogen to liquefied gas is 200 to 300, and the feeding volume airspeed of the liquefied gas is 0.6 to 0.8 h-1。
6. The process of claim 1 wherein the hydrogenation catalyst is a supported hydrogenation catalyst or a bulk hydrogenation catalyst.
7. The process of claim 6 wherein the supported hydrogenation catalyst comprises a support selected from the group consisting of Al and a supported active metal component2O3SiO-containing2Al of (2)2O3、TiO2Molecular sieve-containing Al2O3And one or more of the group consisting of active carbon, wherein the active metal component is selected from noble metals or non-noble metals, the noble metals comprise one or more of Pt, Pd and Re, and the non-noble metals are selected from one or more of W, Mo, Ni and Co.
8. The process of claim 7, wherein the noble metal is present in an amount of 0.1 to 2.0 wt.%, based on the weight of the catalyst; the content of non-noble metal is 5wt% -35 wt% calculated by metal oxide.
9. The process of claim 6, wherein the bulk hydrogenation catalyst comprises three metal components of Mo, W and Ni, wherein W, Ni exists in the form of composite oxide: nixWyOzZ ═ x +3y, Mo is present in the form of oxides of MoO3(ii) a Composite oxide NixWyOzThe atomic mol ratio of the x and the y is 1: 8-8: 1, and the composite oxide Ni isxWyOzAnd oxide MoO3The weight ratio of (A) to (B) is 1: 10-10: 1; composite oxide Ni in bulk hydrogenation catalystxWyOzAnd oxide MoO3The total weight content of the components is 40-100%.
10. The method according to claim 9, wherein the complex oxide Ni is NixWyOzThe atomic molar ratio of the x to the y is 1: 4-4: 1; composite oxide NixWyOzAnd oxide MoO3The weight ratio of (A) to (B) is 1: 5-5: 1; composite oxide Ni in bulk hydrogenation catalystxWyOzAnd oxide MoO3The total weight content of the composition is 50 to 80 percent.
11. The method as set forth in claim 1, wherein decane as said heat transfer medium has a purity of 90wt% or more.
Priority Applications (1)
| Application Number | Priority Date | Filing Date | Title |
|---|---|---|---|
| CN201610991238.4A CN108069815B (en) | 2016-11-11 | 2016-11-11 | Method for preparing high-purity butane |
Applications Claiming Priority (1)
| Application Number | Priority Date | Filing Date | Title |
|---|---|---|---|
| CN201610991238.4A CN108069815B (en) | 2016-11-11 | 2016-11-11 | Method for preparing high-purity butane |
Publications (2)
| Publication Number | Publication Date |
|---|---|
| CN108069815A CN108069815A (en) | 2018-05-25 |
| CN108069815B true CN108069815B (en) | 2020-10-16 |
Family
ID=62154749
Family Applications (1)
| Application Number | Title | Priority Date | Filing Date |
|---|---|---|---|
| CN201610991238.4A Active CN108069815B (en) | 2016-11-11 | 2016-11-11 | Method for preparing high-purity butane |
Country Status (1)
| Country | Link |
|---|---|
| CN (1) | CN108069815B (en) |
Families Citing this family (1)
| Publication number | Priority date | Publication date | Assignee | Title |
|---|---|---|---|---|
| US11680216B2 (en) * | 2019-01-29 | 2023-06-20 | Lanzatech, Inc. | Production of bio-based liquefied petroleum gas |
Citations (4)
| Publication number | Priority date | Publication date | Assignee | Title |
|---|---|---|---|---|
| CN101081998A (en) * | 2006-05-31 | 2007-12-05 | 中国石油化工股份有限公司 | Hydrogenation method for C4 distillation cut |
| CN101113126A (en) * | 2006-07-26 | 2008-01-30 | 李莉 | Olefin-containing lighter hydrocarbons catalytic hydrogenation method |
| CN102311783A (en) * | 2010-07-07 | 2012-01-11 | 中国石油化工股份有限公司 | Liquefied petroleum gas-coke gasoline hydrogenation combined process method |
| CN102452879A (en) * | 2010-10-14 | 2012-05-16 | 中国石油化工股份有限公司 | Method for producing ethylene cracking material through hydrogenation of liquefied petroleum gas |
-
2016
- 2016-11-11 CN CN201610991238.4A patent/CN108069815B/en active Active
Patent Citations (4)
| Publication number | Priority date | Publication date | Assignee | Title |
|---|---|---|---|---|
| CN101081998A (en) * | 2006-05-31 | 2007-12-05 | 中国石油化工股份有限公司 | Hydrogenation method for C4 distillation cut |
| CN101113126A (en) * | 2006-07-26 | 2008-01-30 | 李莉 | Olefin-containing lighter hydrocarbons catalytic hydrogenation method |
| CN102311783A (en) * | 2010-07-07 | 2012-01-11 | 中国石油化工股份有限公司 | Liquefied petroleum gas-coke gasoline hydrogenation combined process method |
| CN102452879A (en) * | 2010-10-14 | 2012-05-16 | 中国石油化工股份有限公司 | Method for producing ethylene cracking material through hydrogenation of liquefied petroleum gas |
Non-Patent Citations (3)
| Title |
|---|
| 《列管式反应器温度控制方法》;肖建良等;《广州化工》;20130131;第41卷(第2期);第120-121、154页 * |
| 《液化气加氢反应动力学研究》;方向晨等;《石油化工》;20130415;第42卷(第4期);第399-403页 * |
| 液化气加氢热力学的研究》;艾抚宾等;《石油化工高等学校学报》;20130215;第26卷(第1期);第29-32、44页 * |
Also Published As
| Publication number | Publication date |
|---|---|
| CN108069815A (en) | 2018-05-25 |
Similar Documents
| Publication | Publication Date | Title |
|---|---|---|
| Li et al. | Coke formation on Pt–Sn/Al 2 O 3 catalyst in propane dehydrogenation: coke characterization and kinetic study | |
| TW201825449A (en) | Dehydrogenation of lpg or ngl and flexible utilization of the olefins thus obtained | |
| Ouyang et al. | Catalytic conversion of bio-ethanol to ethylene over La-modified HZSM-5 catalysts in a bioreactor | |
| CN103962062B (en) | The packing method of isothermal reactor catalyst | |
| Mazoyer et al. | Insights into the deactivation mechanism of supported tungsten hydride on alumina (WH/Al2O3) catalyst for the direct conversion of ethylene to propylene | |
| CN109280561A (en) | A kind of method of naphtha or the low-temperature catalyzed reaction propylene co-production aromatic hydrocarbons processed of lighter hydrocarbons | |
| Scholz et al. | Ethylene to 2‐Butene in a Continuous Gas Phase Reaction using SILP‐Type Cationic Nickel Catalysts | |
| CN104107718B (en) | Catalyst for manufacturing olefin by low-carbon alkane dehydrogenation and preparation method thereof | |
| CN102443430B (en) | Method for preparing ethylene cracking material by hydrogenising liquefied petroleum gas | |
| CN102933527B (en) | Process for co-oligomerization of olefins | |
| TW200829535A (en) | Propylene production | |
| CN105622310B (en) | A kind of method for producing low-carbon alkene and aromatic hydrocarbons | |
| CN108069815B (en) | Method for preparing high-purity butane | |
| CN107537461A (en) | The catalyst of dehydrogenating low-carbon alkane and its application | |
| CN104549260B (en) | Catalyst for preparing olefin through dehydrogenation of isobutene and preparation method of catalyst | |
| Pisarenko et al. | Modeling the process of purifying ethylene from acetylene hydrocarbons over palladium nanocatalysts | |
| Tong et al. | Kinetic studies on the dimerization of isobutene with Ni/Al2O3 as a catalyst for reactive distillation process | |
| CN105722953A (en) | Method of converting a coal to chemicals | |
| CN103864564A (en) | Technique for processing methanol-to-propylene by-products | |
| CN102452879B (en) | For the method for industrial C 4 fraction producing ethylene cracking material through hydrogenation | |
| CN102311787B (en) | Method of hydrogenating liquefied petroleum gas to prepare ethylene cracking feed | |
| CN108779399A (en) | Method for producing antiknock component and catalyst | |
| JP6405857B2 (en) | Method for producing conjugated diene | |
| CN107721791A (en) | The preparation system and preparation method of a kind of preparing propylene from methanol | |
| CN105733668B (en) | A kind of Hydrotreatment method for coal tar |
Legal Events
| Date | Code | Title | Description |
|---|---|---|---|
| PB01 | Publication | ||
| PB01 | Publication | ||
| SE01 | Entry into force of request for substantive examination | ||
| SE01 | Entry into force of request for substantive examination | ||
| GR01 | Patent grant | ||
| GR01 | Patent grant |