TWI487561B - Separation process and superfractionation system - Google Patents

Description

Separation method and super fractionation system

The embodiments described herein relate to the manufacture of linear alpha olefins, and more particularly to linear alpha olefin manufacturing processes having reduced energy consumption.

Linear alpha olefins can be made by selective oligomerization of ethylene. Generally, the oligomerization is carried out in the presence of a catalyst such as an alkylated metal catalyst. Long residence times are used to make hydrocarbon chains of different lengths. When each ethylene molecule is added, the olefinic double bond remains at the alpha position. The ethylene oligomerization reaction produces a wide range of linear alpha olefin products. A large amount of fractionation is required to separate alpha olefins having different carbon numbers.

U.S. Patent No. 3,776,974 is directed to a process for the manufacture of linear monoolefins useful in detergents by disproportionation of linear alpha olefins with 2-butene to produce linear beta olefins having lower and higher carbon numbers. The resulting lower carbon number beta olefin fraction is subjected to a disproportionation reaction to produce 2-butene and a linear internal monoolefin.

U.S. Patent No. 6,727,396 to U.S. Patent No. 6,727,396, issued to U.S. Pat. No. 6, 727,396, the entire disclosure of which is incorporated herein by reference. A linear alpha olefin having a carbon number undergoes a metathesis reaction. The reaction product includes ethylene and a linear internal olefin (LIO) having a carbon number greater than the first number of carbon atoms. Specifically, a feed linear alpha olefin having a carbon number of n is subjected to a metathesis reaction with a second feed linear alpha olefin to form ethylene and a linear internal olefin having a carbon number of 2n-2. The resulting linear internal olefin can then be isomerized to produce a linear alpha olefin. One of the ones In an embodiment, the autometathesis reaction of 1-butene is carried out to produce ethylene and 3-hexene under conditions in which the framework and/or double bond isomerization can be minimized or eliminated and in the presence of a metathesis catalyst. The reaction product of an alkene. The 3-hexene is then isomerized to 1-hexene. In the autometathesis reaction, the catalyst and reaction conditions should be catalysts and reaction conditions which minimize the isomerization of the 1-butene starting material. The catalyst used in the autometathesis reaction may be a carrier catalyst or a carrier-free catalyst and the entire catalyst system has a catalyst having a minimum amount of acidic and basic sites. Typical isomerization catalysts include alkali metal oxides or promoted zeolites.

Superfractionation is known to be useful for separating two components having similar boiling points. In general, superfractionation involves a large number of theoretical trays of distillation columns and high reflux ratios and therefore requires a large amount of input energy to provide reboiling and condensation loads.

Systems and methods for providing reduced energy consumption in the manufacture of linear alpha olefins involving superfractionation can be useful.

One embodiment comprises the steps of: (a) separating a mixed butene stream comprising 1-butene and 2-butene into a top 1-butene stream and a bottoms residue in a butene distillation column. a butene stream separating a portion of the bottoms residue 2-butene stream to form a butene reboiler stream which is heated in a reboiler and gasified and returned to the butene distillation column, b) subjecting at least a portion of the overhead 1-butene stream of (a) to a catalytic metathesis reaction to produce an effluent comprising 3-hexene, (c) isomerizing 3-hexene of (b) to produce a a mixed hexene stream of hexene, 2-hexene and 3-hexene, (d) separating the mixed hexene stream in a hexene fractionation column to form a 1-hexene vapor overhead stream condensed in a cooler And a bottoms stream comprising 2-hexene and 3-hexene, and (e) heating (a) butene reboiling by heat obtained by condensing the (1-) 1-hexene vapor overhead stream Logistics.

Another embodiment is a system for producing 1-hexene from butene, comprising a butene distillation column, a metathesis reactor, a metathesis recovery zone, a hexene isomerization reactor, a hexene fractionation column, and Condenser. The butene distillation column is configured to separate 1-butene from 2-butene and has a butene reboiler including a first heat source coupled thereto. The metathesis reactor converts 1-butene to 3-hexene and ethylene. The metathesis recovery zone separates 3-hexene from ethylene. The hexene isomerization reactor is configured to isomerize 3-hexene to 1-hexene and 2-hexene, and the hexene distillation column is configured to separate 1-hexene vapor from 2-hexene and 3-hexene. The condenser condenses 1-hexene vapor. The first heat source of the butene reboiler is at least a portion of the heat of condensation of the 1-hexene vapor.

A further embodiment comprises the following separation method: obtaining a hydrocarbon mixture comprising a first hydrocarbon, a second hydrocarbon, and a higher boiling hydrocarbon, the second hydrocarbon having a boiling point at atmospheric pressure higher than a boiling point of the first hydrocarbon at atmospheric pressure 0.3-10 ° C. The hydrocarbon mixture is separated into a first feed stream and a second feed stream. The first feed stream is supplied to the first intermediate stage of the first distillation column and the second feed stream is supplied to the second intermediate stage of the second distillation column. The first distillation column is operated at a top pressure higher than the second distillation column. The first distillate containing a majority of the first hydrocarbon is removed from the top of the first distillation column, at least partially condensed, and supplied to the second distillation column at a third intermediate stage above the second intermediate stage. The higher boiling hydrocarbons are removed from the bottom of the first distillation column. The first hydrocarbon is removed from the top of the second distillation column and the second hydrocarbon is removed from the bottom of the second distillation column. At least a portion of the second hydrocarbon removed from the bottom of the second distillation column is heated with the first distillate and returned to the second distillation column.

Another embodiment is an ultrafractionation system comprising a first distillation zone and a second distillation zone. The first distillation zone includes a first distillation column, a first condenser having a first coolant source, and a first reboiler having a first heat source. The first distillation column has a first feed inlet, a first distillate outlet, and a first bottoms residue outlet configured to receive a first portion of fresh feed. The second distillation zone includes a second distillation column configured to operate at a lower pressure than the first distillation column, a second condenser, and a second reboiler having a second heat source. The second distillation column has a second feed inlet configured to receive a second portion of fresh feed, a third feed disposed above the second feed inlet and configured to receive at least a portion of the distillate of the first distillation zone A feed inlet, a second distillate outlet, and a second bottoms residue outlet. For the second reboil The second heat source of the device includes a first coolant source for the first condenser.

100‧‧‧The whole system

102‧‧‧1-butene/2-butene splitter

104‧‧‧ reboiler

106‧‧‧hexene shunt

108‧‧‧Reflow tank

110‧‧‧1-hexene vapor stream

112‧‧‧ Logistics

114‧‧‧ reboiled logistics

116‧‧‧ Butene bottom residue logistics

118‧‧‧ heat exchanger

120‧‧‧Cooling water flow

122‧‧‧Return Logistics

124‧‧‧1-hexene product logistics

200‧‧‧ twin tower system

202‧‧‧First Distillation Tower

204‧‧‧Second distillation tower

206‧‧‧ Logistics

211‧‧‧ Logistics

212‧‧‧ Logistics

213‧‧‧ Logistics

214‧‧‧ Logistics

215‧‧‧ Logistics

216‧‧‧Condenser

217‧‧‧ Logistics

218‧‧‧Return Logistics

220‧‧‧ distillate logistics

222‧‧‧ Logistics

223‧‧‧1-hexene product logistics

224‧‧‧ Logistics

225‧‧‧Return Logistics

226‧‧‧ Logistics

228‧‧‧ reboiler

232‧‧‧Logistics

234‧‧‧ Logistics

250‧‧‧The whole system

252‧‧‧First Tower

254‧‧‧Second Tower

256‧‧‧Incoming logistics

258‧‧‧ Logistics

259‧‧‧ reboiler

260‧‧‧ Logistics

261‧‧‧ Logistics

262‧‧‧ Logistics (distillate)

263‧‧‧ Logistics

264‧‧‧ Logistics

265‧‧‧ reboiled logistics

266‧‧‧Condenser

268‧‧‧Return Logistics

270‧‧‧Second Tower Feed Logistics

272‧‧‧ Logistics

273‧‧‧Return Logistics

274‧‧‧Condenser (logistics)

275‧‧‧1-hexene product logistics

276‧‧‧ Logistics

278‧‧‧ reboiler

280‧‧‧Remaining bottom residue logistics

282‧‧‧ Logistics

284‧‧‧C6 Clearance Logistics

286‧‧‧Remaining logistics

250'‧‧‧The whole system

252'‧‧‧First Tower

254'‧‧‧Second Tower (fractionator)

256'‧‧‧ feed

258'‧‧‧First tower feed

259'‧‧‧ reboiler

260'‧‧‧Second tower feed

262'‧‧‧ Distillate

266'‧‧‧Condenser

268'‧‧‧Return Logistics

270'‧‧‧Second Tower Feed Logistics

272'‧‧‧ distillate

273'‧‧‧Liquid Logistics

275'‧‧‧Vapor flow

290‧‧‧Compressor (fractionator)

292‧‧‧Separator (hexene isomerization reactor)

291‧‧‧ bottom residue logistics

294‧‧‧ fractionator

296‧‧‧isomerization reactor

295‧‧‧ side cut

302‧‧‧hexene isomerization and purification zone

304‧‧‧ Butene isomerization and purification zone

305‧‧‧butene feed logistics

306‧‧‧Automatic metathesis zone

307‧‧‧ Logistics

308‧‧‧Reactor effluent logistics

309‧‧‧ fractionator

310‧‧‧Top Logistics

311‧‧‧Recycling Logistics

312‧‧ ‧ double decomposition recycling logistics

313‧‧‧Incoming logistics

314‧‧‧ Logistics

316‧‧‧ Logistics

318‧‧‧ fractionation section

320‧‧‧ Logistics

322‧‧‧Clear logistics

324‧‧ ‧ complex decomposition product logistics

325‧‧‧ Logistics

326‧‧‧ fractionation section

328‧‧‧1-hexene product logistics

329‧‧‧C6 Clearance Logistics

330‧‧‧ heat exchanger

331‧‧‧Reactor

332‧‧‧ heat exchanger

333‧‧‧reactor

334‧‧‧ heat exchanger

335‧‧‧Reactor

336‧‧‧ heat exchanger

337‧‧‧ Logistics

338‧‧‧ heat exchanger

339‧‧‧C7+ Clearance Logistics

340‧‧‧ heat exchanger

341‧‧‧Return Logistics

342‧‧‧ furnace

344‧‧‧ furnace

346‧‧‧ furnace

1 is a schematic illustration of a portion of a system for producing 1-hexene according to one embodiment; FIG. 2 is a schematic diagram of a conventional dual pressure superfractionation system; and FIG. 3 is a graph showing the comparison of energy consumption of certain disclosed embodiments with conventional systems. Figure 4 is a schematic illustration of a dual pressure superfractionation system of the embodiments described herein; Figure 5 shows a variation of the system of Figure 4, wherein an open loop thermal pump is used; and Figure 6 schematically illustrates the conventional fractionation of linear alpha olefins. - Isomerization system; Figure 7 schematically shows a novel fractionation-isomerization system for obtaining linear alpha olefins; and Figure 8 is a flow diagram showing the overall process for producing 1-hexene according to certain embodiments.

The embodiments described herein provide a linear alpha olefin manufacturing process that reduces the energy used in manufacturing. In one embodiment, the overhead condenser of the hexene separation column is used as a reboiler for the butene separation column. This reduces the overall heating requirements of the system. In another embodiment, which may be used alone or in combination with the first embodiment, a two column fractionation system is used for one or both of the butene and hexene purification columns. The feed is separately supplied to two separate integrated columns operating at two different pressures, and the heat exchange takes place in a manner that achieves lower cost production.

The term "superfractionation" as used herein refers to the separation carried out by distillation of two or more components having a boiling point difference of between 0.3 and 10 °C. As used herein, "catalytic metathesis reaction" refers to the disproportionation of two other different olefins between two olefins.

There are three general process stages in the manufacture of hexene from the metathesis reaction involving the automatic metathesis of butene. First, 1-butene is produced from a C4 stream (e.g., a residue 2 stream). Thereafter, 1-butene undergoes self-disproportionation or auto-metathesis to form ethylene and 3-hexene. Finally, 3-hexene is isomerized to 1-hexene, followed by purification of the 1-hexene. Each of these process stages is described in further detail below.

In the production of 1-butene, the residual liquid 2 is supplied to a combined distillation/butene isomerization system. Residual liquid 2 is a C4 stream comprising butane, butene-1 and butene-2 with significantly reduced amounts of isobutylene and butadiene. The distillation product typically contains an overhead stream comprising 1-butene in a purity of at least 90% by weight. Due to the low relative volatility between the isomers of n-butene, a very large number of distillation trays having extremely high reflux ratios are required to achieve separation. Under these demands, the distillation column requires a large amount of energy to supply reboil and condensation loads.

In the case of 1-butene self-disproportionation or auto-metathesis to form ethylene and 3-hexene, 1-butene is reacted on a metathesis catalyst and the product of the reaction is separated in a fractionation system. There are many fractionation options in the fractionator sequence because the product includes several different carbon number products from C2 to C7. The requirements for the fractionation system are (1) separation of C2/C3 products for recovery of valuable olefins, (2) recycling of C4/C5 olefins to achieve high process yields, and (3) separation of major 3-hexenes. Logistics is the feed to the final process stage. Non-limiting examples of suitable auto-metathesis catalysts are Group VI B or Group VII B metal oxides such as tungsten oxide, molybdenum oxide, or cerium oxide.

Purification was carried out in a second "super-fractionator" in the final step of 1-hexene production (i.e., isomerization of 3-hexene to 1-hexene and purification of 1-hexene). This distillation is even more difficult than distillation for 1-butene fractionation. A large number of trays and extremely high reflux ratios are required, which results in considerable energy consumption. In summary, the energy required for butene separation and hexene separation indicates that the energy consumption of the entire process is quite large. Several methods and systems described herein reduce energy usage in the manufacture of linear alpha olefins such as 1-hexene.

One method of reducing the energy consumption in the manufacture of 1-hexene from butene is to use a hexene-shunt top tray vapor to heat the bottom tray liquid of the butene-shunt. This eliminates the heating energy consumption of the n-butene separation system and the overhead cooling requirements of the hexene separation system.

Both the butene-shunt and the hexene-shunt require a higher concentration of alpha-olefin in the overhead condenser liquid as product. In many cases, the butene splitter requires 1-butene in the product stream to be greater than 90%, and sometimes greater than 95%. Hexene splitters usually require a product The 1-hexene in the stream is greater than 98%. The relative volatility of internal olefins (2-butene, 2-hexene and 3-hexene) and alpha olefins (1-butene and 1-hexene) in the butene and hexene columns is so low that fractionation Two towers are required with a large number of stages and reflux. The relative volatility is the vapor pressure ratio of the individual components at the same temperature and represents the relative tendency of fractionation separation. The higher the relative volatility, the easier the separation.

In conventional systems, the butene splitter is designed to operate at the lowest possible pressure to increase the relative volatility between 1-butene and 2-butene and thereby enhance the desired separation. Increasing the relative volatility reduces the required backflow and heating/cooling load levels. However, the temperature necessary to cool the condenser liquid limits the butene column pressure. The very low butene pressure causes the overhead temperature to be outside the cooling water temperature range and therefore requires refrigeration and additional process energy to provide refrigeration. For this application, the (etc.) butene column can generally be designed to be about 70 psig.

For the same reasons as above, the hexene shunt of the conventional system was designed in the same manner as the butene splitter, i.e., the separation was enhanced by increasing the relative volatility using low column pressure. Since hexene has a higher boiling point than butene, the hexene column can be operated at atmospheric pressure or below atmospheric pressure and still uses cooling water in the condenser. This column is typically operated at a pressure between -10 and 5 psig.

In one embodiment described herein, the hexene splitter is operated at a higher pressure to enable the overhead vapor stream to be used to heat the butene splitter reboiler stream. Operating the (equal) hexene splitter at higher pressures at a constant reflux ratio produces more theoretical orders because the relative volatility decreases with increasing absolute pressure. Increasing the top pressure of the hexene splitter from 0 psig to 15 psig, as indicated by the computer simulations described in Table 1 below, increases the theoretical separation order from 103 to 125. This caused a 21% increase. Therefore, higher pressure towers have higher capital costs due to the increased number of stages.

Alternatively, if the same number of stages is used at an overhead pressure of 15 psig, the higher pressure column will require a higher reflux ratio to achieve the same degree of separation. This will increase the tower diameter (resulting in increased capital costs) and energy usage (caused by increased reboiling and condensing loads).

During operation of the hexene column to provide a reboiling load to the butene column, the butene column pressure should preferably be greater than 70 psig, or in the range of from about 70 psig to about 100 psig, to avoid the use of refrigeration in the condenser. This produces a reboiling temperature of about 139 °F in the C4 column. By operating the hexene column at high pressure, the overhead condenser (the coldest temperature in the hexene column) can be raised to the point where a hexene condenser can be used as the reboiler for the C4 column. Although this would increase the fractionation requirements (stage, reflux, or both) of the hexene tower, surprisingly, increasing the pressure of the hexene column to allow the use of a condenser to provide a reboiling load on the butene column would result in net energy savings. .

In order to minimize the heat exchange between the bottom tray liquid of the 1-butene/2-butene splitter and the 1-hexene/2-hexene and 3-hexene top tray vapor, the pressure of the hexene tower must be increased. Up to the "high pressure" range (15 psig) described in Table 1. Operating at a lower pressure in the hexene column reduces the minimum approach of the heat exchanger, as shown in Table 2 below:

Operating the hexene column at atmospheric pressure produces a vapor temperature at the top of the tray that does not effectively heat the butane bottoms tray liquid but optimally fractionates the hexene. The use of high hexene tower pressure is counterintuitive, so more trays are needed to achieve the desired separation. However, by using a high hexene column pressure, the butene vapor can be used to heat the butene reboiler stream. This will save energy usage throughout the system.

Figure 1 shows a system with the above thermal integration. Specify the entire system as 100. The system comprises a 1-butene/2-butene splitter 102 and a reboiler 104 with 1-hexene/2-hexene and 3-hexene splitter 106 and a reflux tank 108. 1-Hexene vapor stream 110 The top of the olefinic splitter 106 is removed and used as a heating stream in the reboiler 104, which reduces the temperature of the stream 110 passing through the reboiler 104 and typically condenses the stream 110 to some extent. Stream 112 is removed from the bottom of butene splitter 102 and is separated into reboil stream 114 and butene bottoms stream 116. Stream 114 is heated in reboiler 104 and then returned to the lower end of butene splitter 102.

After the heat is provided in the heat exchanger 104, if desired, the 1-hexene vapor stream 110 is further cooled in the heat exchanger 118 using the cooling water stream 120 until it is completely condensed. The cooled 1-hexene stream 110 is passed to a reflux tank 108. A portion of the effluent from tank 108 is returned to reflux 16 as reflux stream 122 and the remainder is removed as 1-hexene product stream 124. Under the above circumstances, the condenser load required for the hexene tower 106 is greater than the reboiler duty of the butene column 102. In an alternative design in which the butene component is more dilute or the product 1-butene stream and the 1-butene stream to be metathesized are taken out, the load of the butene splitter reboiler should be greater than that of the hexene splitter condenser. load. In such cases, an additional reboiler of stream 114 may be substituted for the additional cooler of stream 110 to balance the reboiler duty of the butene splitter.

The degree of cooling water/steam reduction in hexene and butene splitters depends on the column specification and throughput. The thermal integration illustrated in connection with Figure 1 can result in a significant reduction in cooling water (CW) and heated steam (LPS). Tables 3 and 4 below compare the energy usage of the process before and after the integration of the hexene and butene columns for 50 KMTA 1-hexene production capacity.

Notes: 1. In autometathesis, column 1 is depropanized (DEC3) and column 2 is depentene (DEC5). (i) The *-DEC3 condenser has a refrigerant of -5 °C. (ii) The **-DEC5 reboiler has high pressure steam. (iii) A 2500 KW CW exchange occurs on the DEC3 to reduce the refrigerant load. The values are displayed separately. 2. In C6 isomerization, column 1 has a higher pressure and column 2 has a lower pressure than column 1. 3. Energy Integration - (a) In the C6 isomerization system, the column 1 condenser is exchanged with the column 2 reboiler. Both cases are shown as net value = 0. Part of the two-column system (b) exchange of the column 2 condenser in the C6 isomerization with the reboiler in the C4 isomerization. (Note - in the first table - the values reflect the load before the exchange.)

As can be seen from these tables, the CW requirement is reduced from 76,231 KW to 48,771 KW or 36% by the exchange between the butene reboiler and the hexene condenser, and the low pressure steam heating load (LPS) is reduced from 68,301 KW to 39,901 KW. Or reduced by 42%.

To achieve high manufacturing capabilities, multiple 1-butene/2-butene splitters and multiple 1-hexene/2-hexene and 3-hexene splitters can be used. In this case, the combined fluid shown as stream 112 of the butene-shunt can be exchanged with the combined fluid shown as stream 110 of the hexene-divider. In some cases, an additional tempering cooling water exchanger may be required.

Another technique for increasing process efficiency separates the feed to the butene splitter and/or the feed to the hexene splitter into two separate columns and operates the two columns at different pressures in the superfractionation process. When used in hexene separation, this feature of the disclosed embodiment reduces the energy consumption of the hexene system by about 33% at a comparable number of trays and operating pressure. Such savings are generally applicable to superfractionation systems that can separate hydrocarbons having low relative volatility. For hexene systems, the advantages achieved by this embodiment are combined with the advantages that can be achieved by the integrated heat exchange system described above.

In reboiling and condensing operations, the superfractionator consumes a lot of energy. Several published studies show various options for reducing the energy cost of a superfractionator. Known systems include: (a) a dual pressure column in which the column is divided into two different columns. The first column is operated at a higher pressure than the second column and the condenser of the first column is used to reboil the second column, and (b) the thermal pump is used. In this case, the overhead vapor is compressed to a higher pressure. The compression energy heats the vapor to a higher temperature. The higher pressure/higher temperature vapor is subsequently condensed (at higher temperatures) to provide a reboiling load to the column. This choice does not save energy by itself, but allows operation under conditions that save energy. By reducing the pressure of the columns, the relative volatility of the hydrocarbons can be increased. This allows for the use of fewer stages, or less reflow, or both to achieve a given separation. However, operating at lower pressures can lower the temperature of the column and may require a different condenser refrigeration load than the cooling water. This will increase energy consumption. By operating the thermal pump, both the lower temperature load and the reboiling load can be provided by the compression power that does not result in increased energy due to additional refrigeration.

System (a) (two column system) has been used in various separation column sequences to, for example, separate low carbon olefins from their paraffin analogs. If a single column is designed to facilitate operation, the 100% condenser load is in the top of the column and the 100% reboiling load is in the bottom of the column. This represents the maximum temperature difference of the tower. By using two column fractionation at two different pressures, the columns can be operated at different reflux ratios and the load can be optimized in terms of both total energy and energy levels. The energy level can be defined as the cooling temperature in many cases, where the higher the degree of cooling in the refrigeration system, the more energy is required. This integration differs from the use of a tower side reboiler or a tower side intercooler to regulate the reflux and temperature distribution within a single column to optimize utility.

According to the embodiments described herein, when 1-hexene is separated from 2-hexene and 3-hexene in a two-column system with split feed, the difference in top pressure between the two distillation column columns is typically 4-40 psi. , or 6-20 psi, or 8-15 psi. When separating other components having similar boiling points using the two-column system described herein, the pressure differential across the top of the two distillation column columns is typically in the range of 4-40 psi, or 6-20 psi, or 8-15 psi. The optimum choice of pressure difference is determined by the desired temperature difference of the heat exchange and the final size (and cost) of the exchanger.

The key to effectively using system (a) is to balance the operation of each of the two columns to match the condensation load of the first column to the desired reboiling load of the second column. In some cases, an imbalance is allowed and this difference can be compensated for by installing a smaller external heating reboiler that compensates for the difference, or a smaller cooling condenser. This is usually used to control the system. Another possibility is to adjust the fractionation load in each column to achieve equilibrium. The usual way of balancing this operation is by making A portion of the vapor of the first column can be bypassed to the second column, or a portion of the liquid of the second column can be transferred to the first column to transfer the load from one column to the other. In a first conventional option, the column feed is delivered to a first or higher pressure column. The amount of condensation load of the first column is controlled by manipulating the vapor bypass from the first column to the second column. This bypass transfers the load from the condensing portion of one column to the reboiling load of the other column, so the condenser of column 1 should be matched to the reboiler duty of column 2. The savings achieved by this type of system are primarily in energy level (temperature) rather than total energy savings. In addition, for a tower with a large number of stages, dividing the tower into two separate towers will reduce costs compared to a single pole tower.

A comparison between a conventional single column system and a conventional twin column system having a vapor bypass between the first column and the second column is provided below. Cases A and B represent a single tower system with different capacities and number of trays. Case E uses a lower pressure drop tray (for cases A and B, tower dP = 20 kPa versus 35 kPa). Cases C and D represent a two-column system with similar capacity but different number of trays. A single hexene system typically has a conventional distillation unit of the order of 100+ (typically 150+). The rating is numbered from the top. The feed enters at level 75. The product 1-hexene was removed as a distillate. The bottom residue of 2-hexene and 3-hexene is recycled to a hexene isomerization reactor (not shown) wherein 2-hexene and 3-hexene are isomerized to give 1-hexene A mixture of alkenes, 2-hexene and 3-hexene and returning the effluent to the column. In this way, 3-hexene is converted to 1-hexene. A small amount of debris was removed from the system. The tower specification is given in Table 5. Cases A and B operate at different flow rates and pressures and different stages and produce 1-hexene with slightly different purity.

The conventional two-pressure distillation separation of 1-hexene was carried out in two towers each having a stage of 100 (providing a total of 200 stages). A schematic of the two-column system is shown in Figure 2 and is typically designated 200. The system includes a first distillation column 202 and a second distillation column 204. The feed containing about 8% 1-hexene and the remainder comprising mainly 2-hexene and 3-hexene and a small amount of C7+ material enters the first column 202 in stream 206 at stage 40. The bottom product of the column in stream 211 is primarily C7+, which is purged in stream 213, leaving the remainder in stream 215 to the bottom of column 202. In stream 212, the distillate product of the first column has 35% 1-hexene. Vapor side fraction from the first Stage 81 of a tower 202 is withdrawn and enters stream 217 to balance the exchange load. If the first column 202 is integrated with the isomerization reaction system, the liquid side draw is withdrawn from stage 87 and passed to stream 214 for supply to the isomerization reaction system. The recycle stream comprising mainly 2-hexene and 3-hexene is mixed with fresh 3-hexene feed (not shown) and passed through an isomerization reactor (not shown) to produce from 2 and 3-hexene. A certain 1-hexene is returned to the first column as feed 206.

The distillate of first column 202 in stream 212 is condensed in condenser 216 and then separated into distillate stream 220 and reflux stream 218 entering the top of first column 202. Distillate stream 220 enters second column 204 at level 54. The vapor side cut in stream 217 enters second column 204 at stage 90. Stream 217 enters at a lower level than stream 220 because it is extracted at a lower stage and thus contains less 1-hexene. The product 1-hexene is removed as a distillate from second column 204 to stream 222, condensed, and separated into 1-hexene product stream 223 and reflux stream 225. The bottoms product of second column 204 in stream 224 is reboiled in reboiler 228 and a portion is returned to second column 204 in stream 226. Another portion is combined with the liquid side draw of the first column in stream 214 to form stream 232 which is recycled to the hexene isomerization reactor. A small amount of scavenger can be removed from stream 234 if desired.

The second column 204 operates at a lower pressure than the first column 202. This enables the condenser 216 of the first column 202 to operate at a higher temperature than the reboiler 228 of the second column 204, facilitating heat exchange therebetween. Thus, thermal integration is achieved by the condenser connecting the first column 202 with the reboiler of the second column 204. The individual column specifications are given in Table 6. Two different designs are provided in which the total fractionation level can vary. The first case uses 100 trays (200 in total) in each of the two towers, while the second case uses 150 trays (300 in total) in each of the two towers.

The energy consumption of single tower scenarios A-B and E and twin tower scenario C-D is shown in Tables 7 and 8 below.

When the results are tabulated based on energy usage, it is clear that the total energy usage of the two systems at the same distillation stage is similar. This is shown in the graph of Figure 3. The energy usage of both the single column and the double column vapor bypass decreases along a similar curve. There is a slight difference depending on the operating pressure, where the specific energy of the lower pressure single column case is about 10% lower. It can be seen that Case A has a high recovery (strict specification) compared to Case B. This has no significant effect on energy. However, Case B and Case E have similar recovery rates but Case D has a significantly lower tray dP. This will affect relative volatility.

The advantage of a twin column system with vapor bypass is not the energy consumed per ton of product but the ability to increase the reboil temperature of the first column. In this case, the reboiling temperature was raised from 99 ° C in the case of a single column to 124 ° C in the case of a double column. The two-tower vapor bypass system is best suited for refrigeration systems where the reboiler requires refrigeration. In many cases, this increase in temperature can shift the load from a low temperature requiring refrigeration compression to a higher temperature level, which requires less refrigeration or may be reboiled by steam or lower cost energy. The transfer of the reboiling temperature clarifies the principle when the illustrated case does not use refrigeration. It should be noted that in many cases with a large number of trays (>100), the tower is designed as two series towers to reduce costs by reducing the number of foundations and structures. If the tower is to be constructed as two series columns, the dual pressure configuration is preferred.

Other embodiments disclosed herein are dual pressure twin column systems with feed separation and heat integration and corresponding methods. These systems have reduced energy consumption compared to the conventional dual pressure twin tower systems described above. The system and method are shown schematically in Figure 4. Specify the entire system as 250. Although the drawings and description relate to the separation of 1-hexene from 2-hexene and 3-hexene, the scope of the examples also includes the manufacture of other types of components having similar boiling points.

Distillation separation of 1-hexene with 2-hexene and 3-hexene is carried out in columns 252 and 254. In the illustrated embodiment, column 252 has 80 stages and column 254 has 70 stages. Therefore, the total number of stages (150) is the same as in the case of one of the above single towers. It should be noted that other levels can actually be used.

a feed split of the separation system, wherein a portion of the feed stream 256 is used as stream 258 Level 25 enters the first tower 252. The other portion enters the second column 254 as a stream 260 at level 60. The bottom product of column 252 in stream 261 is primarily C7 and most of the product is purged in stream 263. Reboiler 259 heats reboiler stream 265 to return it to first column 252. The distillate of first column 252 in stream 262 has 40% 1-hexene. The liquid side fraction containing 2-hexene and 3-hexene is taken as the stream 264 from the 74th stage of the first column 252 and supplied to the isomerization reaction system (not shown). The recycle stream comprising primarily 2-hexene and 3-hexene is mixed with fresh 3-hexene feed (not shown) through an isomerization reactor (not shown) from 2-hexene and 3 Hexene produces a certain 1-hexene and it returns to the first column as feed 256.

The distillate 262 of the first column 252 is condensed in the condenser 266 and the condensate is separated into a reflux stream 268 that returns to the top of the first column 252 and a second column feed stream 270 that enters the second column 254 at the 40th stage, The 40th level is at the feed point of the second portion of the fresh feed entering the 60th stage. It should be noted that the side cuts and feed locations that are not those described herein may be used.

The product 1-hexene is removed from the second column 254 as a distillate to stream 272. Distillate 272 is condensed in condenser 274 and separated into reflux stream 273 and 1-hexene product stream 275. The bottoms product of second column 254 in stream 274 is partially reboiled in reboiler 278. The reboiled stream is separated into a stream 276 that returns to the bottom of the second column 254 and a remaining bottoms stream 280. The bottoms stream 280 is combined with the liquid side draw of the first column in stream 264 as stream 282. A portion of stream 282 is separated as C6 purge stream 284 and the remaining stream 286 is recycled to the hexene isomerization reactor (not shown).

The second column 254 operates at a lower pressure than the first column 252. This allows the condenser of the first column to be at a higher temperature than the reboiler of the second column, facilitating heat exchange therebetween. The loads are balanced by the split of the feed to each column. The relative load of reboiler 278 and condenser 266 can be balanced by delivering more or less fresh feed to column 2. Thus, thermal integration is achieved by connecting the condenser of the first column to the reboiler of the second column.

Other benefits of the process of Figure 4 can be achieved if a thermal pump of recyclable heat and cooling fluid is incorporated between the reboiler 259 of the second column 252 and the condenser 274 of the first column 254. In a In one embodiment, an open loop heat pump utilizing the distillate of the first column 254 can be used, as shown in FIG. In this embodiment, many of the system components are the same, wherein the entire system is designated 250', and the feed 256' is divided into a first column feed 258' entering the first column 252' and a second entering the second column 254'. The tower feeds 260'. Distillate 262' is removed from first column 252', condensed in condenser 266' and separated into reflux stream 268' and second column feed stream 270'. The difference is that the distillate 272' of the second column 254' is compressed in the compressor 290 and used as the heating fluid in the reboiler 259'. Stream 272' is then separated in separator 292 into a vapor stream 275' that is removed as product, and a liquid stream 273' that serves as a reflux stream for fractionator 254'. In this embodiment, the compressor pressurizes the hot vapor to a pressure that produces a desired reboiling temperature.

The computerized simulation tower specifications are given in Table 9. In the table, case F represents a two-column feed splitter. In this example, the pressure of the first split feed column is greatly increased compared to a single column system. The top pressure of the second tower is roughly equivalent to the top pressure of a single tower. If used in combination with this embodiment, this facilitates an equivalent exchange between the hexene condenser and the hexene deoiler.

Table 10 gives an overview of the simulation results. The energy usage of the two-column feed split system was reduced to 1.44 KW/KTA product hexene-1. This is also shown in Figure 3 as "Double Tower Feed Diversion". This value can be compared to 2.13 of a single column system with the same number of trays and comparable tray pressure drop and specifications. This is reduced by 32%. The energy consumption is still 23% higher (1.91 vs. 1.44) when operating a single column with the same low pressure drop tray count (higher tray cost) and low 1-hexene recovery. In addition, this energy consumption is comparable to the energy consumption of a conventional dual pressure twin tower system requiring twice the number of trays. This represents a substantial capital advantage of the disclosed embodiment system. The first column of the two column system operates at a higher pressure than the single column system (or second column). Therefore, most of the fractionation takes place at a higher pressure, which results in a smaller diameter of the column for the part and thus saves capital costs.

Surprisingly, the configuration shown in Figure 4 can significantly reduce the amount of energy used to separate a two-column system containing a mixture having a boiling point very close to the material. In some applications, a split feed column is utilized with a single column in which portions of the split feed have different compositions. For example, it is known to use a split feed column after vapor-liquid separation. In this case, the vapor is supplied at a higher point in the column and the liquid system is supplied at a lower portion of the column to utilize the equilibrium flash separation. In contrast to the established fractionation practice for finding the optimum feed location, the same feed combination with a similar phase (vapor or liquid) is supplied at two separate fractionation columns and in two different columns operating at different pressures. Things. The optimum feed location is typically defined by the point at which a similar composition appears in the column. However, significant energy savings can be achieved by using the principles of the disclosed embodiments as a means of balancing the operation of the dual column-double pressure system.

In the split feed embodiment, the pressure level of each column can be raised or lowered to optimize energy usage. In addition to achieving absolute fractional energy savings, additional energy savings can be achieved in, for example, a refrigeration system with a conventional two-column system.

In yet another embodiment, the side take position of stream 264 of first column 252 in Figure 4 can be selected based on the details of the system design to minimize the 1-hexene content of the stream. As described above, since the relative volatility between 1-hexene and 2-hexene and 3-hexene is close, high purity 1- Separation of hexene (>99 mol.%) requires a large number of stages and column reflux. As a result, the hexene fractionation step is an energy intensive step. The hexene fractionator removes 1-hexene from the top of the column and removes 2-hexene and 3-hexene from the bottoms. The 2-hexene and 3-hexene bottoms are recycled to the hexene isomerization reactor which forms more 1-hexene. The hexene isomerization reactor effluent is combined with the fresh 3-hexene feed and returned to the hexene fractionator. The concentration of 1-hexene is balanced by the equilibrium at the exit temperature of any given isomerization reactor. Operation at high temperatures is limited by fouling and operation at lower temperatures is subject to unfavorable equalization. The single pass conversion of the isomerization reactor is set by uniformly setting the 1-hexene inlet concentration of the recycle stream to the outlet concentration of the reactor outlet.

Those who are constructing a system for the manufacture of linear alpha olefins (e.g., 1-hexene) are expected to design a low concentration of 1-hexene that can be obtained in the bottoms of the fractionator (1 mol.% is shown in Mode 1 below). And, in turn, a fractionation system that increases the single pass conversion of the downstream hexene isomerization reactor. The lowest possible concentration of 1-hexene in the column system is at the bottom of the column. The single column scheme seems reasonable because the low concentration of 1-hexene in the bottoms of the fractionator reduces the amount of hexene recycle. Figure 6 shows a conventional fractionator 290. A portion of the bottoms bottoms stream 291 is recycled to the hexene isomerization reactor 292 and then returned to the fractionator 290. The design of Figure 6 (mode 1 below, bottoms with low concentrations of 1-hexene) also produces high 1-hexene recovery (the total amount of 1-hexene produced with fresh feed hexene) Ratio), due to the minimal loss of 1-hexene during the removal process.

The solution of Figure 7 can be used in conjunction with the embodiment of Figure 4 and can also be incorporated into the embodiment of Figure 2. In the embodiment depicted in Figure 7 (Mode 2), the concentration of 1-hexene in the hexene side fraction removed is higher than that present in the bottom of the column and is supplied to the hexene isomerization reactor. More specifically, fractionator 294 has a side cut 295 that is passed to isomerization reactor 296. One non-limiting example of the 1-hexene concentration of the side cut is 2.5 mole % 1-hexene. The greater concentration of 1-hexene in the reactor feed reduced the hexene reactor single pass conversion and increased the hexene isomerization reactor recycle rate by 2.7%. This increase will increase the energy requirements for the isomerization reactor. The hexene scavenger in Figure 7 (Mode 2) is still removed from the bottom of the column under the same 1-hexene bottoms specification (bottom residue) ()). Mode 2 increases the recycle rate of the hexene isomerization reactor by 25%, thereby increasing the utility of the isomerization system including gasification and preheating, and thus the reflux flow rate inside the hexene tower is never significantly reduced significantly. . As a result, for the same 1-hexene recovery, there was a net decrease in the column and isomerization energy consumption and reflux ratio when mode 2 was used. The reason for this net decrease in reflux flow of the column is the reduction in energy required to meet the slightly loose hexene isomerization recycle specification.

When Mode 2 was used instead of Mode 1, the reduced energy usage is shown in Table 11 below.

As shown in Table 11, when Mode 2 was used, the total energy usage was reduced from 93.59 MM Kcal/hr to 88.58 MM Kcal/hr with a reduction in energy of about 5%.

Certain embodiments of the auto-metathesis process described in U.S. Patent No. 6,727,396, which includes isomerization of butenes for the conversion of internal butenes and internal hexene olefins to external alpha olefins, respectively. Reactor and hexene isomerization reactor. In addition, the process uses an automated metathesis reactor system. Each of the three reaction systems is a vapor phase fixed bed reaction system. Each system requires gasification of the feed to the reactor (since the feed is liquid when supplied to the unit). In addition, for each reactor, the effluent temperature is significantly higher than the feed temperature and/or the temperature of the fractionation column after the reactor. These temperatures are required because the higher temperatures in the isomerization step favor the desired alpha olefin and the reaction kinetics require a higher temperature for the autometathesis reaction. In conventional systems, the effluent from the reactor can be used to vaporize and preheat the feed to some extent and the remaining preheating heat can be provided by the flame heater.

8 shows a flow chart of a method of making 1-hexene in which additional heat exchangers 330, 332, and 334 are added to reduce the capital cost of furnaces 342, 344, and 346. The system shown in Figure 8 includes hexene isomerization and purification zone 302, butene isomerization and purification zone 304, and auto-metathesis zone 306. The butene feed stream 305 is combined with recycle stream 311 to form stream 307. Stream 307 is heated in heat exchanger 330, further heated in heat exchanger 336, sent to furnace 342 and subsequently isomerized in reactor 331. The reactor effluent stream 308 is fractionated in a fractionator 309 and the overhead stream 310 is separated into a reflux stream 341 and a metathesis feed stream 313 that is passed to a metathesis zone 306. The bottoms bottoms stream 331 is recycled and may be partially purged in stream 337. In metathesis zone 306, stream 313 is combined with metathesis recycle stream 312 to form stream 314. Stream 314 is heated in heat exchanger 334, further heated in heat exchanger 340, sent to furnace 346, and metathesized in reactor 335. The reactor effluent in stream 316 is fractionated in fractionation section 318. The byproducts are removed in stream 320 and purge stream 322 is extracted. C4 and C5 are recycled in stream 312. The metathesis product stream 324 is passed to the hexene isomerization and purification zone 302.

In zone 302, stream 324 is heated in heat exchanger 332, heated in heat exchanger 338, sent to furnace 344, and subsequently isomerized in reactor 333. The reactor effluent in stream 325 is fractionated in fractionation section 326 to form 1-hexene product stream 328, C6 purge stream 329, and C7+ purge stream 339. By adding individual heat exchangers 330, 332, 334 to preheat the upstream reactor feed from the butene isomerization reactor 331, the hexene isomerization reactor 333, and the autometathesis reactor 335, respectively, also in the furnace upstream of each of the reactors. Reduced consumption. Most of the required heat load is present in gasifiers 330, 332, and 334 that are located upstream of process/process exchangers 336, 338, and 340, and such process/process exchangers 336, 338 and 340 are located directly upstream of furnaces 342, 344 and 346, respectively. Compared to previously known systems that do not include heat exchangers 330, 332, and 334, the flow chart depicted in Figure 8 can reduce furnace capital costs by more than 80%.

Heat exchangers 330, 332, and 334 are typically gasifiers. Such heat exchangers may utilize waste heat from another source, if applicable, if desired. For example, in an ethylene system, this can be quench water. In addition, in any integrated steam system, maximizing the use of "pass-out" steam increases the overall efficiency of the steam system. In order not to waste energy, the flame heater requires additional heat recovery (after the process load is warmed up). The larger flame load of the heater requires a larger amount of additional heat, resulting in an increase in the capital cost of the system. Adding equipment to reduce equipment costs is counterintuitive. However, in this case, a low cost gasifier can be added to reduce the size of the higher cost flame heater, thereby saving money and reducing energy consumption.

It will be appreciated that various of the features and functions disclosed above, as well as other various features and functions, or alternatives thereof, may be desirably combined into many other different systems or applications. At the same time, those skilled in the art can make various alternatives, modifications, variations or improvements of the present invention, and are intended to be included in the following claims.

250‧‧‧The whole system

252‧‧‧First Tower

254‧‧‧Second Tower

256‧‧‧Incoming logistics

258‧‧‧ Logistics

259‧‧‧ reboiler

260‧‧‧ Logistics

261‧‧‧ Logistics

262‧‧‧ Logistics (distillate)

263‧‧‧ Logistics

264‧‧‧ Logistics

265‧‧‧ reboiled logistics

266‧‧‧Condenser

268‧‧‧Return Logistics

270‧‧‧Second Tower Feed Logistics

272‧‧‧ Logistics

273‧‧‧Return Logistics

274‧‧‧Condenser (logistics)

275‧‧‧1-hexene product logistics

276‧‧‧ Logistics

278‧‧‧ reboiler

280‧‧‧Remaining bottom residue logistics

282‧‧‧ Logistics

284‧‧‧C6 Clearance Logistics

286‧‧‧Remaining logistics

Claims (10)

A separation method comprising: (a) providing a fresh feed comprising 1-hexene, 2-hexene and 3-hexene, (b) separating the fresh feed into a first feed stream and a second feed a feed stream, (c) supplying the first feed stream to a first intermediate stage of a first distillation column operating at a first overhead pressure, (d) supplying the second feed stream to a second distillation a second intermediate stage of the column, the second distillation column is operated at a second column pressure lower than the pressure of the first column, (e) removing 1-hex from the top of the first distillation column a first distillate of the alkene, at least partially condensing the first distillate using a first condenser, and at a third intermediate stage between the top and the second intermediate stage of the second distillation column An at least partially condensed first distillate is supplied to the second distillation column, (f) a higher boiling hydrocarbon is removed from the bottom of the first distillation column, and (g) is from the top of the second distillation column Removing the 1-hexene, (h) removing the second bottom product from the bottom of the second distillation column, and (i) using the second reboiler at a temperature lower than the first condenser temperature to perform the first a condenser And heat exchange between the second reboiler, thereby heating at least a portion of the second bottoms product and returning the heated portion to the second distillation column. The method of claim 1, wherein the second distillation column is operated at an average pressure that is at least 7 psi below the average operating pressure of the first distillation column. The method of claim 1, further comprising (j) removing the first liquid as a side draw from a fourth intermediate stage between the bottom of the first distillation column and the first intermediate stage. The method of claim 3, wherein the first portion of the first liquid removed as a side cut is removed. The method of claim 4, wherein the first portion of the first liquid removed as a side fraction is removed and the second portion of the first liquid removed as a side fraction is isomerized. The method of claim 5, wherein the second portion of the first liquid removed as the side fraction is isomerized in the isomerization reactor, and wherein the first liquid removed as the side fraction can be adjusted by adjusting The flow rate is used to control the amount of energy used in the isomerization reactor. The method of claim 1, wherein at the same 1-hexene production rate, the method results in a savings of at least about 10% energy compared to a system in which a single distillation column is used. The method of claim 1, wherein at the same 1-hexene production rate, the method results in a savings of at least about 20% energy compared to a system in which a single distillation column is used. A superfractionation system comprising: a first distillation zone comprising a first distillation column, a first condenser, and a first reboiler, the first distillation column having a first portion configured to receive fresh feed a first feed inlet, a first distillate outlet, and a first bottoms outlet, the fresh feed comprising 1-hexene, 2-hexene and 3-hexene; and a second distillation zone comprising a structured Forming a second distillation column, a second condenser, and a second reboiler operating at a lower pressure than the first distillation column, the second distillation column having a first configuration configured to receive the fresh feed a second portion of the second feed inlet, a third feed inlet, a second distillate outlet located above the second feed inlet and configured to receive at least a portion of the distillate from the first distillate outlet, And a second bottom outlet; wherein the temperature of the second reboiler is lower than the temperature of the first condenser, and wherein the first condenser and the second reboiler exchange heat to heat at least a portion thereof The second bottom outlet removes the bottom product, and the heated portion returns to the first Distillation tower. The superfractionation system of claim 9 further comprising a first side draw outlet positioned on the first distillation column, wherein the first side draw outlet is configured to remove the liquid to be supplied to the isomerization reactor Side stream.