CA2387169A1 - Extractive distillation - Google Patents
Extractive distillation Download PDFInfo
- Publication number
- CA2387169A1 CA2387169A1 CA 2387169 CA2387169A CA2387169A1 CA 2387169 A1 CA2387169 A1 CA 2387169A1 CA 2387169 CA2387169 CA 2387169 CA 2387169 A CA2387169 A CA 2387169A CA 2387169 A1 CA2387169 A1 CA 2387169A1
- Authority
- CA
- Canada
- Prior art keywords
- column
- process according
- stream
- weight
- distillation column
- 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.)
- Abandoned
Links
Classifications
-
- B—PERFORMING OPERATIONS; TRANSPORTING
- B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
- B01D—SEPARATION
- B01D3/00—Distillation or related exchange processes in which liquids are contacted with gaseous media, e.g. stripping
- B01D3/34—Distillation or related exchange processes in which liquids are contacted with gaseous media, e.g. stripping with one or more auxiliary substances
- B01D3/40—Extractive distillation
-
- B—PERFORMING OPERATIONS; TRANSPORTING
- B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
- B01D—SEPARATION
- B01D3/00—Distillation or related exchange processes in which liquids are contacted with gaseous media, e.g. stripping
- B01D3/14—Fractional distillation or use of a fractionation or rectification column
- B01D3/143—Fractional distillation or use of a fractionation or rectification column by two or more of a fractionation, separation or rectification step
Landscapes
- Chemical & Material Sciences (AREA)
- Chemical Kinetics & Catalysis (AREA)
- Organic Low-Molecular-Weight Compounds And Preparation Thereof (AREA)
Abstract
The present invention provides a flexible extractive distillation process to separate one or more C4-10 olefins from a solvent or diluent comprising one or more C4-10 alkanes. The process involves the extractive distillation of the mixture using a lactam in the absence of an additional extractive agent such as a sulfolane or glycol. The process is useful in recovering monomer from solvent /diluent in solution or slurry polymerizations.
Description
FIELD OF THE INVENTION
The present invention relates to a process for the extractive separation of a mixture of components used in the production of polyethylene. The components are mixed with a lactam to separate the streams in a simple and convenient manner.
BACKGROUND OF THE INVENTION
United States Patent 5,085,740 issued Feb. 4, 1992 to Lee et al.
teaches an extractive distillation process for separating close boiling components. Typically the components comprise an alkene and a close boiling alkane. According to the patent the binary mixture is combined with a lactam, a sulfolane and a glycol. The resulting mixture is then passed through a separation (distillation) column to separate an overhead stream rich in alkane and a bottoms stream comprising extraction solvent and alkene. The bottom stream is then passed through a further separation column to separate an overhead stream rich in alkenes and a bottom stream rich in extraction solvent. The present invention has eliminated the use of the essential sulfolane and glycol as taught by the reference.
The present invention seeks to provide a simple efficient process for separating a stream of components from the solution polymerization of 3 o ethylene. Further the present invention seeks to provide an extractive distillation process for vapor recovery has the flexibility to generate a high purity comonomer stream (i.e. 1-butene, 1-hexene, or 1-octene) and polymerization solvent stream without using additional distillation columns or changes in the process configuration to accommodate alternative comonomer. Additionally, the recovered polymerization solvent/
M:\Trevor\TT Specs\9240-Cda(01-18).doc 2 comonomer mixture may be directly recycled to the reaction process without further separation.
SUMMARY OF THE INVENTION
The present invention provides a process for the extractive separation of a mixture of one or more C4_1o alkenes from a mixture of one or more C4_1o alkanes comprising feeding said mixture to a counter current 1o extractive distillation column in which the extractive solvent comprises from 100 to 90 weight % of one or more polar C4_8 lactams and from 0 to weight % of water and separating the mixture into at least one overhead low boiling stream and a higher boiling bottoms stream.
DETAILED DESCRIPTION
Figure 1 is a schematic diagram of the process of the present invention.
Figure 2 shown the effect of entrainer: feed ratio in the extraction column on stream purity at an overhead hexane withdrawal rate of 4500 kg/hr and a reflux ratio of 1 as calculated in example 1.
Figure 3 shows the impact of reflux ratio used in the extraction column (103) on the purity of the hexane rich stream at a 10:1 entrainer:
feed ratio as calculated in example 1, as well as the NMP concentration in 3 o the hexane rich stream.
Figure 4 shows the effect of hexane withdrawal rate (stream 7) and entrainer: feed ratio on a hexane rich stream concentration as calculated in example 1.
As used in this specification the term reflux ratio means the weight ratio of refluxlvapor distillate.
M:\Trevo~\TT Specs\9240-Cda(01-18).doc 3 In the solution polymerization process monomer and comonomer are dissolved in a solvent and subjected to polymerization conditions.
Solution polymerizations may be conducted at temperatures in, but not limited to, the range of 105° C to 200° C and especially in the range of 130° C to 180° C. The polymerization process may be conducted in a reactor system such as in a tubular reactor or multi-reactor system. The 1o pressures used in the process are those known for solution polymerization processes, for example, pressures in the range of about 4-20 MPa. The pressure and temperature are controlled so that the polymer formed remains in solution.
Optionally, small amounts of hydrogen, for example 1-40 parts per million by weight, based on the total solution fed to the reactor may be added to one or more of the feed streams of the reactor system in order to improve control of the melt index and/or molecular weight distribution and thus aid in the production of a more desirable product, as is disclosed in Canadian Patent 703,704.
The polymerization process may be used to prepare copolymers of ethylene and higher alpha-olefins having densities in the range of, for example, about 0.900-0.970 g/cm3 and especially 0.910-0.930 glcm3.
3 o Such polymers may have a melt index as measured by the method of ASTM D-1238, condition E, in the range of, for example, about 0.1-200, and especially in the range of about 0.3 - 45 dg/min. Such a melt index tends to indicate a higher molecular weight of the resulting polymer. The polymers may be manufactured with narrow or broad molecular weight distribution. For example, the polymers may have a stress exponent, a M:\Trevor\TT Specs\9240-Cda(01-18).doc 4 measure of molecular weight distribution, in the range of about 1.1-2.5 and especially in the range of about 1.3-2Ø
Stress exponent is determined by measuring the throughput of a melt indexer at two stresses (21608 and 64808 loading) using the procedures of the ASTM melt index test method, and the following formula:
1o Stress exponent = 1/0.477 X log (wt. of polymer extruded with 64808 wt.)/(wt. of polymer extruded with 2160 g wt.).
Stress exponent values of less than about 1.40 indicate narrow molecular weight distribution while values above about 1.70 indicate broad molecular weight distribution.
The solution passing from the polymerization reactor is normally treated to deactivate any catalyst remaining in the solution. A variety of catalyst deactivators are known, examples of which include but not limited to fatty acids, alkaline earth metal salts of aliphatic carboxylic acids and alcohols. The hydrocarbon solvent used for the deactivator is preferably the same as the solvent used in the polymerization process. If a different solvent is used, it must be compatible with the solvent used in the polymerization mixture and not cause adverse effects on the solvent so recovery system associated with the polymerization process.
The solvent may then be flashed off from the polymer, which subsequently may be extruded into water and cut into pellets or other suitable comminuted shapes. The recovered polymer may then be treated with saturated steam at atmospheric pressure to, for example, reduce the amount of volatile materials and improve polymer color. The treatment M:\Trevo~\TT Specs\9240-Cda(01-18).doc 5 may be carried out for about 1 to 16 hours, following which the polymer may be dried and cooled with a stream of air for 1 to 4 hours.
As a result after the polymerization of the monomers there is a stream comprising a mixture of one or more C4_1o alkenes and a mixture of one or more C4_1o alkanes. Typically the stream comprises at least one C4_ 1o alkene, at least one C2_3 alkene, one or more C1_1o alkane, and oligomers of said alkenes. The streams recovered from the flashing and the steam treatment of the polymer may be combined and treated in accordance with the present invention.
The separation process of the present invention will now be described in conjunction with figure 1.
Figure 1 shows a three column configuration in which column 1, the first column, is a fractional distillation column used to separate the low boiling components (i.e. CO, C02, H2, methane, ethylene, ethane), from the high boiling grease (e.g. products of partial polymerization). Column 1 can be trayed or packed, structured or random.
With respect to butene copolymer production, the composition of the feed stream to column 100 is a function of the desired polymer properties in terms of density, melt index, stress exponent, and physical 3o processing properties of the polymer. Typically, this composition includes the following components and the weight fraction range in which they can appear: polymerization solvent (PS) at 50-99%, ethylene at 0.1-10%, comonomer (i.e. 1-butene typically with 0 to 20, preferably 0-10, weight of 2-butene) at 0-50%, comonomer isomers at 0-50%, inert hydrocarbon (e.g. C1_4 alkanes such as methane, ethane and the like) at 0-20%, trace M:\TrevorlTT Specs\9240-Cda(01-18).doc quantities of hydrogen, CO and C02, preferably only H2 and a small amount, typically less than 1 weight %, preferably from 0.3 to 0.6 weight of oligomers of the monomers (grease).
Feed enters near the middle of column 100, where the temperature can be controlled by a feed heater or cooler (not shown). At the top of the column, overhead vapor is refluxed with a condenser unit 101. The io condenser 101 can be a partial or total condenser, depending on the cooling medium available. The condensed liquid is returned to the column as reflux while the vapor distillate, shown as stream 2, is withdrawn for further processing in downstream unit operations to provide a high purity ethylene stream for recycle back to the polymerization reactors. The reflux ratio was found to be an efFective parameter to control the recovery of ethylene in the vapor distillate. The reflux ratio (weight ratio of reflux/vapor distillate) could vary over the range of 0.1-25:1. A reflux ratio of 15:1 to 20:1, preferably from 16:1 to 18:1, most preferably 17:1 was effective to achieve a 94%weight ethylene recovery in stream 2.
At the bottom of column 100, some of the liquid product withdrawn can be heated in a reboiler, unit 102, and returned to the column above the reboiler. The remainder, stream 3, is withdrawn to downstream unit operations for further processing or placed in storage vessels. The composition of this stream ranges from 0.1 to 50%, preferably 0.1 to 15, most preferably 0.1 to 10 weight% grease, with the remainder being polymerization solvent.
The operating pressure in column 100 can vary over the range of 50 to 1500, preferably from 300 to 1000 kPag (kilopascals gage), most M:\Trevor\TT Specs\9240-Cda(01-18).doc 7 preferably from 500 kPag to 800 kPag at the top, depending on the cooling medium available in the condenser 101. The operating temperature at the top of the column can vary from 10° C to 100° C, preferably from 50° C to 90° C, most preferably from 75° C to 80° C depending on the cooling medium available in the condenser. The operating temperature profile of 78° C at the condenser and 144° C at the reboiler was used at the so corresponding 650 kPag column pressure.
At an appropriate location below the condenser 101, in column 100, a side draw, stream 4, is taken which consist of a mixture of comonomer (e.g. 1-butene, 1-hexene or 1-octene, typically not less than 80, preferably not lest than 90, most preferably not less than 95 weight % of such alkene and the balance isomers of the alkene for example not less than 80 weight 20 % of 1-butene and from 0 up to 20 weight % of 2-butene) and polymerization solvent (e.g. n-pentane, 2-methyl pentane, 3-methyl pentane, n-hexane, cyclohexane, 2-methyl hexane, 3-methyl hexane 2,2-dimethyl butane, 2,3-dimethyl butane, cyclopentane and mixtures thereof).
In one embodiment the polymer solvent is a mixture of alkanes comprising not less than 55 weight % of 2-methyl pentane. This stream can be either liquid or vapor or mixed phase. The composition of the stream, can vary 3o from 0.1 to 50 weight % of comonomer (and isomers thereof e.g. 2-butene) with the remainder being essentially polymer solvent. A portion of this stream, shown as stream 5, is sent to extractive distillation column 103. The remainder, shown as stream 6, is recycled back to the polymerization reactors.
M:\Trevor\T'i' Specs49240-Cda(01-18).doc Stream 5 enters the extractive distillation column 103 near the middle and can be either pre-heated or cooled (not shown). The entrainer solvent (ES), shown as stream 9 enters the column at a point about a1/5th of the height of the column, measured from the top. The location of the entry point was found to be an effective means to control the concentration of entrainer solvent in the overhead distillate product.
Suitable entrainer solvents include polar C4~ lactams including N-methyl pyrrolidone, N-ethyl pyrrolidone, and N-propyl pyrrolidone, preferably N-methyl pyrrolidone (NMP).
The entrainer solvent flows in a downward direction, making contact with the upward flowing feed. This countercurrent flow path promotes intimate contact between the entrainer solvent and the comonomer/polymer solvent mixture. With 1-butene comonomer, this component will concentrate in the overhead distillate, stream 7, with the purity achieved in the range of 60-95% weight.
The flow rate of entrainer solvent used is a function of the capacity of the entrainer solvent and its selectivity. The weight ratio of entrainer:feed required can vary from 0.01-20:1, preferably 0.1 to 10, depending on the composition of the feed stream and the desired purity of 3o the product streams. For a composition of 10% weight 1-butene in the feed stream 5 to the extraction column 103, a ratio of entrainer solvent to feed rate of 0.1:1 by weight was found to be effective in producing a high purity comonomer (e.g. 1-butene) distillate stream 7 of 95% weight concentration. This comonomer stream may then be recycled to the polymerization reactors.
M:lTrevor\TT Specs\9240-Cda(01-18).doc The vapor overhead in column 103 is condensed by the exchanger unit 104. The reflux ratio was found to be an effective operating parameter to control the composition of the overhead distillate, stream 7, particularly the level of entrainer solvent. The reflux ratio (weight ratio of reflux/liquid distillate) can vary over the range of 0.1-20:1. At the operating temperatures and pressures over a 35 ideal staged column, no detectable level of entrainer solvent was observed in the composition of the distillate stream 7 when the reflux ratio of 2:1 was used.
Any suitable operating pressure range in the extraction column 103 can be used to effect the separation. The pressure may be from 1200 to 1600 kPag, preferably from 1300 to 1500 kPag, most preferably about 1400 kPag at the top of the column was used. A pump (not shown) to boost the feed pressure of stream 5 may be required to match the entry pressure in column 103. The operating temperature profile of less than 100° C, typically from 75° C to 100° C preferably from 85° C to 95° C at the top and less than 190° C, typically from 170° C to 190° C, preferably, from 175° C to 185 ° C, most preferably about 178° C at the bottom of the column was used at the corresponding 1400 kPag column pressure.
For butene copolymers, at an appropriate height, between the overhead distillate and the entry location of the entrainer solvent, a side draw (not shown) can be withdrawn to remove the trans-2-butene isomer.
The side draw can be either a liquid or vapor and can be sent to downstream unit operations for further processing or to storage vessels.
The bottom stream 10 is concentrated with the entrainer solvent and the polymerization solvent. The composition may be from 80 to 95 M:\Trevor\TT Specs\9240-Cda(01-18).doc weight %, preferably 85 to 95 weight most preferably about 90% weight polymer solvent and 20 to 5 weight %, preferably 15 to 5 weight %, most preferably about 10% weight entrainer solvent. A portion of this stream, stream 11, is heated through a reboiler 105, and then returned to the bottom of the extraction column 103. The remainder, shown as stream 12, is fed to the third or down stream stripper column 107, In Figure 1, a 1o process heat exchanger 106, to recover the heat from stream 17 from the bottom of the stripper column 107 is shown. This heat integration configuration is optional and serves to illustrate the flexibility of the process to reduce energy operating costs.
Stream 13 exiting heat exchanger 106 contains approximately 90 weight % of polymer solvent and 10 weight % of entrainer solvent.
Separation of polymer solvent and entrainer solvent occurs by fractional distillation in column 107 with polymer solvent recovered in the overhead distillate stream 14 and the entrainer solvent at the bottom of stripper column 107 in stream 15. Stream 13 enters column 107 near the top, at approximately 1/5t" of the height of the column, measured from the top.
The feed location was found to be effective in the control of entrainer solvent that carries over into stream 14 and thereby reducing the purity of 3 o the polymer solvent.
The vapor overhead from column 107 is condensed in condenser unit 108, where reflux is returned to the column and the liquid distillate stream 14 is withdrawn. The reflux ratio was found to be an effective operating parameter to control the composition of the overhead distillate, stream 14, particularly the level of entrainer solvent. A reflux ratio (weight M:\Trevor\TT Specs19240-Cda(01-18).doc 11 ratio of reflux/liquid distillate) of 0.5:1 was found to be sufficient but may range from 0.1:1 to 10:1. A purity in excess of 99.9 weight % polymer solvent was obtained in stream 14. The quantity of stream 14 generated is sufficient for use in the catalyst preparation area of the polymerization process with any excess recycled directly to the polymerization reactors.
A purity of 99.6 weight % entrainer solvent in the bottom stream 15 1o was obtained. A portion of the bottom product stream 15 is heated through a reboiler, 109, and then returned to the bottom of the column, shown as stream 16. The remainder, shown as stream 17 is heat exchanged in the process/process exchanger 106, with the discharge directed to the top of the extraction column 103 to complete the recycle of the entrainer solvent stream. A small purge stream 18, is withdrawn to avoid accumulation of impurities in the entrainer solvent stream. Fresh makeup of entrainer solvent may be added back into the entrainer solvent recycle, shown as stream 8.
Any suitable operating pressure range in the stripper column 107 can be used to effect the separation. The pressure of 50 to 500 kPag, preferably from 130 kPag to 150 kPag , typically 140 to 150 (e.g. 145) kPag at the top of the column was used. The operating temperature 3o profile from 65° C to 90° C, typically from 68° C to 72° C, preferably about 70° C at the top and from 80° C to 110° C, preferably from 85° C to 90° C, typically about 87° C at the bottom of the column was used at the corresponding 145 kPag column top pressure.
For the production of hexene and octene copolymer, where 1-hexene and 1-octene are used respectively as the comonomer, the M:\Trevor\TT Specs\9240-Cda(01-18).doc 12 primary difference is where the polymerization solvent (PS) and the comonomer streams are recovered. In this scenario, a high purity polymerization solvent (PS) stream is recovered in the overhead of the extraction column 103, shown as stream 7. A high purity comonomer stream is recovered in the overhead of the stripper column 107, shown as stream 14. At an appropriate location below the stripper condenser, unit 108, a side draw can be taken to remove the comonomer isomers.
No additional distillation columns or changes in the process configuration are required to accommodate the use of 1-butene, 1-hexene, or 1-octene in the process of the present invention.
For the separation of hexene monomer from the polymerization solvent the extractive distillation column is typically operated at pressures from 50 to 500 kPag and a temperature profile at the top of the column from 70° C to 90° C, preferably from 75° C to 85°
C and a temperature profile at the bottom of the column from 170° C to 200° C, preferably from 180° C to 195° C. In the stripper column (107) the pressure is generally from 50 to 500 kPag and the temperature profile at the top of the column is from 65° C to 90° C, preferably from 70° C to 80°
C, and the temperature profile at the bottom of the column is from 170° C to 200° C, preferably from 180° C to 195° C.
For the separation of octene monomer from the polymerization solvent the extractive distillation column is typically operated at pressures from 50 to 500 kPag and a temperature profile at the top of the column from 65° C to 90° C, preferably from 75° C to 90°
C and a temperature profile at the bottom of the column from 130° C to 160° C, preferably from M:\Trevor\TT Specs\9240-Cda(01-18).doc 13 140° C to 155° C. In the stripper column (107) the pressure is generally from 50 to 500 kPag and the temperature profile at the top of the column is from 125° C to 145° C, preferably from 130° C to 140° C, and the temperature profile at the bottom of the column is from 200° C to 230° C, preferably from 210° C to 230° C.
The present invention will now be illustrated by the following non 1o limiting example. In the example unless otherwise stated the compositions are indicated in weight %.
A computer model developed using PRO/II, Simulation Sciences Inc. was used to model the process of the present invention. The model accurately reflects the solution polymerization processes of NOVA
Chemicals solution polymerization process within less than 3% error.
2o The extractive distillation process was modeled using a commercial process simulator to analyze the separation performance for a number of selected entrainers. In each case, a number of design and operating parameters were varied to determine the impact on product purity and recovery, including: feed tray location, reflux ratio, solvent circulation rate, and number of trays. These steady-state models were used to obtain a comparison of entrainer performance to arrive at a family/class of solvents 3o that show potential as an effective solvent for alkanelalkene separation and to develop a heat and material balance around the extraction process.
In order to develop a steady-state model of the extractive distillation process, thermodynamic methods need to be selected to characterize the vapor-liquid interaction between each component pairing in our system, M:\Trevor\TT Specs\9240-Cda(01-18).doc 14 particularly interaction with the entrainer component. In our study, the following two liquid activity coefficient methods were used:
~ Dortmund modification of UNIFAC
~ Non-Random Two Liquid (NRTL) The following references are suggested for a more detail treatment on the subject of thermodynamic characterization of excess functions using to liquid activity coefficient methods and heats of mixing.
~ The Dortmund modification of UNIFAC {J.Lohmann, et. al., "From UNIFAC to Modified UNIFAC (Dortmund)", Ind. Eng. Chem. Res., 2001, 40, pg. 957-964}.
~ NRTL method ~H.Renon, et. al., "Local Compositions in Thermodynamic Excess Functions For Liquid Mixtures", AIChE
Journal, 1, Vol. 14, 1968, pg. 135-144}.
The Dortmund modification of UNIFAC method provides a quick and reasonably accurate basis with which to perform a preliminary analysis of the effectiveness of various entrainers on product purity and recovery of alkane/alkene systems. Based on these results, the cyclic amide family/class of compounds (i.e. NMP) was selected as possible entrainers for further study.
To improve the level of accuracy in the extraction models, experimental measurement of the vapor-liquid equilibrium between the various binary pairings are needed, particularly for the solute pairings with the entrainer solvent. J. Gmehling, et al, has compiled an extensive collection of vapor-liquid equilibrium data that dates from 1977-present.
The DECHEMA series includes experimental data that covers a wide M:\Trevor\TT Specs\9240-Cda(01-18).doc 15 range of solute and solvent pairings, particularly those with commercial applications in the chemical industries. In addition, heats of mixing for the binaries are also reported along with activity coefficients at infinite dilution.
The latter is used to provide a metrics to rapidly compare and screen alternative entrainers for various separation systems.
Experimental data on vapor-liquid equilibrium of NMP with 1-to hexene and the components in our polymerization solvent has been published by K.Fischer and J.Gmehling (K.Fischer, et al., "Vapor-Liquid Equilibria, Activity Coefficients At Infinite Dilution and Heats of Mixing For Mixtures of N-Methyl Pyrrolidone-2 with C5 or C6 Hydrocarbons and for Hydrocarbon Mixtures", Fluid Phase Equilibria, 119, 1996, pg. 113-130.) that contains VLE data along with heats of mixing. The VLE data was 2o regressed to obtain coefficients for the 8-parameter form of the NRTL
equation for liquid activity coefficients for the various binary pairs. In each case, the regressed equation matches the experimental data over the entire composition range with an average deviation of 0.5%. Maximum deviation of 12.5% was observed at either ends of the curve (i.e. dilute regions).
Experimental heats of mixing data (K.Fischer, 1996) for various binary pairs was also regressed to the 8-parameter Redlich-Kister equation for excess heats of mixing. These parameters were used in the simulation model for the extraction column to provide a more accurate temperature profile across each tray. The maximum deviation between the regressed curve and the experimental data points is less than 1.6%.
M;\Trevor\TT Specs19240-Cda(01-18).doc 16 A computer model was developed using PRO/II, Simulation Sciences Inc., to model the extraction process using the NRTL parameters for k-value prediction and the Redlich-Kister parameters for the heats of mixing for liquid enthalpy prediction. The data is summarized in the three examples below.
EXAMPLES
Example 1 The feed system to be separated is a 50%wt 1-hexene/50% wt n-hexane mixture at a rate of 10,000 kg/hr using NMP as the entrainer solvent. The relative volatility between 1-hexene/n-hexane is merely 1.01 and is reflected in a difference of only 5° C in the normal boiling point. This mixture is a difficult one to separate by fractional distillation alone. The 2o model is used to provide an assessment of the selectivity and capacity of NMP along with the sensitivity on product purity and recovery with variations in entrainer:feed ratio, reflux ratio, and hexane product withdraw rate. Figure 2 shows the impact of entrainer:feed ratio on the purity of the hexane-rich stream and the hexene-rich stream. At an entrainer:feed weight ratio of 10:1, the purity of the hexane-rich stream recovered is 91 wt hexane with the remaining 9% wt being 1-hexene. The purity of the corresponding hexene-rich stream recovered is 84% wt 1-hexene with the remaining 16% wt being hexane. The purity of these two streams is more than sufficient for use in hexene solution copolymerization.
Figure 3 shows the impact of reflux ratio in the extraction column on the purity of the hexane-rich stream recovered, at three different product M:\TrevorlTT Specs\9240-Cda(01-18).doc 17 withdraw rates (i.e. hexane-rich stream at 4000, 4500, and 5000 kglh). In addition, Figure 3 shows the variation of NMP concentration in the hexane-rich stream at the various withdraw rates. As the hexane-rich withdraw rate increases, the concentration of NMP entrained decreases. The ppm levels detected are sufficiently low to enable recycle of the hexane-rich stream to the polymerization reaction.
to Figure 4 shows a 3-dimensional plot of the purity of the hexane-rich stream with variations in entrainer:feed ratio and hexane-rich withdraw rate.
The operating conditions for columns 100,103 and 107were substantially as shown in table 4 in example 3.
Example 2 For butene copolymer mode of operation, a butene/PS mixture is generated from polymerization. Using the extractive distillation process as described in this invention, a polymerization solvent-rich stream is generated (i.e. Stream 14) as overhead from column 107. In addition, a butene-rich stream is generated (i.e. Stream 7) as overhead from column 103. The purity of these two streams is summarized in Table 1. The operating conditions to achieve this separation are provided in Table 2.
M:lTrevor\TT Specs\9240-Cda(01-18).doc 1$
Table 1. Product Purity For Butene Mode Operation @ 0.1:1 Entrainer:Feed Ratio. Feed Rate=100 kglhr Composition Polymerization Butene Comonomer Solvent Recycle Recycle (Stream (Stream 14) 7) Ethylene (wt%) 0 0.8 Polymerization Solvent 99.99 3.3 (wt %) 1-Butene (wt%) 0 95
The present invention relates to a process for the extractive separation of a mixture of components used in the production of polyethylene. The components are mixed with a lactam to separate the streams in a simple and convenient manner.
BACKGROUND OF THE INVENTION
United States Patent 5,085,740 issued Feb. 4, 1992 to Lee et al.
teaches an extractive distillation process for separating close boiling components. Typically the components comprise an alkene and a close boiling alkane. According to the patent the binary mixture is combined with a lactam, a sulfolane and a glycol. The resulting mixture is then passed through a separation (distillation) column to separate an overhead stream rich in alkane and a bottoms stream comprising extraction solvent and alkene. The bottom stream is then passed through a further separation column to separate an overhead stream rich in alkenes and a bottom stream rich in extraction solvent. The present invention has eliminated the use of the essential sulfolane and glycol as taught by the reference.
The present invention seeks to provide a simple efficient process for separating a stream of components from the solution polymerization of 3 o ethylene. Further the present invention seeks to provide an extractive distillation process for vapor recovery has the flexibility to generate a high purity comonomer stream (i.e. 1-butene, 1-hexene, or 1-octene) and polymerization solvent stream without using additional distillation columns or changes in the process configuration to accommodate alternative comonomer. Additionally, the recovered polymerization solvent/
M:\Trevor\TT Specs\9240-Cda(01-18).doc 2 comonomer mixture may be directly recycled to the reaction process without further separation.
SUMMARY OF THE INVENTION
The present invention provides a process for the extractive separation of a mixture of one or more C4_1o alkenes from a mixture of one or more C4_1o alkanes comprising feeding said mixture to a counter current 1o extractive distillation column in which the extractive solvent comprises from 100 to 90 weight % of one or more polar C4_8 lactams and from 0 to weight % of water and separating the mixture into at least one overhead low boiling stream and a higher boiling bottoms stream.
DETAILED DESCRIPTION
Figure 1 is a schematic diagram of the process of the present invention.
Figure 2 shown the effect of entrainer: feed ratio in the extraction column on stream purity at an overhead hexane withdrawal rate of 4500 kg/hr and a reflux ratio of 1 as calculated in example 1.
Figure 3 shows the impact of reflux ratio used in the extraction column (103) on the purity of the hexane rich stream at a 10:1 entrainer:
feed ratio as calculated in example 1, as well as the NMP concentration in 3 o the hexane rich stream.
Figure 4 shows the effect of hexane withdrawal rate (stream 7) and entrainer: feed ratio on a hexane rich stream concentration as calculated in example 1.
As used in this specification the term reflux ratio means the weight ratio of refluxlvapor distillate.
M:\Trevo~\TT Specs\9240-Cda(01-18).doc 3 In the solution polymerization process monomer and comonomer are dissolved in a solvent and subjected to polymerization conditions.
Solution polymerizations may be conducted at temperatures in, but not limited to, the range of 105° C to 200° C and especially in the range of 130° C to 180° C. The polymerization process may be conducted in a reactor system such as in a tubular reactor or multi-reactor system. The 1o pressures used in the process are those known for solution polymerization processes, for example, pressures in the range of about 4-20 MPa. The pressure and temperature are controlled so that the polymer formed remains in solution.
Optionally, small amounts of hydrogen, for example 1-40 parts per million by weight, based on the total solution fed to the reactor may be added to one or more of the feed streams of the reactor system in order to improve control of the melt index and/or molecular weight distribution and thus aid in the production of a more desirable product, as is disclosed in Canadian Patent 703,704.
The polymerization process may be used to prepare copolymers of ethylene and higher alpha-olefins having densities in the range of, for example, about 0.900-0.970 g/cm3 and especially 0.910-0.930 glcm3.
3 o Such polymers may have a melt index as measured by the method of ASTM D-1238, condition E, in the range of, for example, about 0.1-200, and especially in the range of about 0.3 - 45 dg/min. Such a melt index tends to indicate a higher molecular weight of the resulting polymer. The polymers may be manufactured with narrow or broad molecular weight distribution. For example, the polymers may have a stress exponent, a M:\Trevor\TT Specs\9240-Cda(01-18).doc 4 measure of molecular weight distribution, in the range of about 1.1-2.5 and especially in the range of about 1.3-2Ø
Stress exponent is determined by measuring the throughput of a melt indexer at two stresses (21608 and 64808 loading) using the procedures of the ASTM melt index test method, and the following formula:
1o Stress exponent = 1/0.477 X log (wt. of polymer extruded with 64808 wt.)/(wt. of polymer extruded with 2160 g wt.).
Stress exponent values of less than about 1.40 indicate narrow molecular weight distribution while values above about 1.70 indicate broad molecular weight distribution.
The solution passing from the polymerization reactor is normally treated to deactivate any catalyst remaining in the solution. A variety of catalyst deactivators are known, examples of which include but not limited to fatty acids, alkaline earth metal salts of aliphatic carboxylic acids and alcohols. The hydrocarbon solvent used for the deactivator is preferably the same as the solvent used in the polymerization process. If a different solvent is used, it must be compatible with the solvent used in the polymerization mixture and not cause adverse effects on the solvent so recovery system associated with the polymerization process.
The solvent may then be flashed off from the polymer, which subsequently may be extruded into water and cut into pellets or other suitable comminuted shapes. The recovered polymer may then be treated with saturated steam at atmospheric pressure to, for example, reduce the amount of volatile materials and improve polymer color. The treatment M:\Trevo~\TT Specs\9240-Cda(01-18).doc 5 may be carried out for about 1 to 16 hours, following which the polymer may be dried and cooled with a stream of air for 1 to 4 hours.
As a result after the polymerization of the monomers there is a stream comprising a mixture of one or more C4_1o alkenes and a mixture of one or more C4_1o alkanes. Typically the stream comprises at least one C4_ 1o alkene, at least one C2_3 alkene, one or more C1_1o alkane, and oligomers of said alkenes. The streams recovered from the flashing and the steam treatment of the polymer may be combined and treated in accordance with the present invention.
The separation process of the present invention will now be described in conjunction with figure 1.
Figure 1 shows a three column configuration in which column 1, the first column, is a fractional distillation column used to separate the low boiling components (i.e. CO, C02, H2, methane, ethylene, ethane), from the high boiling grease (e.g. products of partial polymerization). Column 1 can be trayed or packed, structured or random.
With respect to butene copolymer production, the composition of the feed stream to column 100 is a function of the desired polymer properties in terms of density, melt index, stress exponent, and physical 3o processing properties of the polymer. Typically, this composition includes the following components and the weight fraction range in which they can appear: polymerization solvent (PS) at 50-99%, ethylene at 0.1-10%, comonomer (i.e. 1-butene typically with 0 to 20, preferably 0-10, weight of 2-butene) at 0-50%, comonomer isomers at 0-50%, inert hydrocarbon (e.g. C1_4 alkanes such as methane, ethane and the like) at 0-20%, trace M:\TrevorlTT Specs\9240-Cda(01-18).doc quantities of hydrogen, CO and C02, preferably only H2 and a small amount, typically less than 1 weight %, preferably from 0.3 to 0.6 weight of oligomers of the monomers (grease).
Feed enters near the middle of column 100, where the temperature can be controlled by a feed heater or cooler (not shown). At the top of the column, overhead vapor is refluxed with a condenser unit 101. The io condenser 101 can be a partial or total condenser, depending on the cooling medium available. The condensed liquid is returned to the column as reflux while the vapor distillate, shown as stream 2, is withdrawn for further processing in downstream unit operations to provide a high purity ethylene stream for recycle back to the polymerization reactors. The reflux ratio was found to be an efFective parameter to control the recovery of ethylene in the vapor distillate. The reflux ratio (weight ratio of reflux/vapor distillate) could vary over the range of 0.1-25:1. A reflux ratio of 15:1 to 20:1, preferably from 16:1 to 18:1, most preferably 17:1 was effective to achieve a 94%weight ethylene recovery in stream 2.
At the bottom of column 100, some of the liquid product withdrawn can be heated in a reboiler, unit 102, and returned to the column above the reboiler. The remainder, stream 3, is withdrawn to downstream unit operations for further processing or placed in storage vessels. The composition of this stream ranges from 0.1 to 50%, preferably 0.1 to 15, most preferably 0.1 to 10 weight% grease, with the remainder being polymerization solvent.
The operating pressure in column 100 can vary over the range of 50 to 1500, preferably from 300 to 1000 kPag (kilopascals gage), most M:\Trevor\TT Specs\9240-Cda(01-18).doc 7 preferably from 500 kPag to 800 kPag at the top, depending on the cooling medium available in the condenser 101. The operating temperature at the top of the column can vary from 10° C to 100° C, preferably from 50° C to 90° C, most preferably from 75° C to 80° C depending on the cooling medium available in the condenser. The operating temperature profile of 78° C at the condenser and 144° C at the reboiler was used at the so corresponding 650 kPag column pressure.
At an appropriate location below the condenser 101, in column 100, a side draw, stream 4, is taken which consist of a mixture of comonomer (e.g. 1-butene, 1-hexene or 1-octene, typically not less than 80, preferably not lest than 90, most preferably not less than 95 weight % of such alkene and the balance isomers of the alkene for example not less than 80 weight 20 % of 1-butene and from 0 up to 20 weight % of 2-butene) and polymerization solvent (e.g. n-pentane, 2-methyl pentane, 3-methyl pentane, n-hexane, cyclohexane, 2-methyl hexane, 3-methyl hexane 2,2-dimethyl butane, 2,3-dimethyl butane, cyclopentane and mixtures thereof).
In one embodiment the polymer solvent is a mixture of alkanes comprising not less than 55 weight % of 2-methyl pentane. This stream can be either liquid or vapor or mixed phase. The composition of the stream, can vary 3o from 0.1 to 50 weight % of comonomer (and isomers thereof e.g. 2-butene) with the remainder being essentially polymer solvent. A portion of this stream, shown as stream 5, is sent to extractive distillation column 103. The remainder, shown as stream 6, is recycled back to the polymerization reactors.
M:\Trevor\T'i' Specs49240-Cda(01-18).doc Stream 5 enters the extractive distillation column 103 near the middle and can be either pre-heated or cooled (not shown). The entrainer solvent (ES), shown as stream 9 enters the column at a point about a1/5th of the height of the column, measured from the top. The location of the entry point was found to be an effective means to control the concentration of entrainer solvent in the overhead distillate product.
Suitable entrainer solvents include polar C4~ lactams including N-methyl pyrrolidone, N-ethyl pyrrolidone, and N-propyl pyrrolidone, preferably N-methyl pyrrolidone (NMP).
The entrainer solvent flows in a downward direction, making contact with the upward flowing feed. This countercurrent flow path promotes intimate contact between the entrainer solvent and the comonomer/polymer solvent mixture. With 1-butene comonomer, this component will concentrate in the overhead distillate, stream 7, with the purity achieved in the range of 60-95% weight.
The flow rate of entrainer solvent used is a function of the capacity of the entrainer solvent and its selectivity. The weight ratio of entrainer:feed required can vary from 0.01-20:1, preferably 0.1 to 10, depending on the composition of the feed stream and the desired purity of 3o the product streams. For a composition of 10% weight 1-butene in the feed stream 5 to the extraction column 103, a ratio of entrainer solvent to feed rate of 0.1:1 by weight was found to be effective in producing a high purity comonomer (e.g. 1-butene) distillate stream 7 of 95% weight concentration. This comonomer stream may then be recycled to the polymerization reactors.
M:lTrevor\TT Specs\9240-Cda(01-18).doc The vapor overhead in column 103 is condensed by the exchanger unit 104. The reflux ratio was found to be an effective operating parameter to control the composition of the overhead distillate, stream 7, particularly the level of entrainer solvent. The reflux ratio (weight ratio of reflux/liquid distillate) can vary over the range of 0.1-20:1. At the operating temperatures and pressures over a 35 ideal staged column, no detectable level of entrainer solvent was observed in the composition of the distillate stream 7 when the reflux ratio of 2:1 was used.
Any suitable operating pressure range in the extraction column 103 can be used to effect the separation. The pressure may be from 1200 to 1600 kPag, preferably from 1300 to 1500 kPag, most preferably about 1400 kPag at the top of the column was used. A pump (not shown) to boost the feed pressure of stream 5 may be required to match the entry pressure in column 103. The operating temperature profile of less than 100° C, typically from 75° C to 100° C preferably from 85° C to 95° C at the top and less than 190° C, typically from 170° C to 190° C, preferably, from 175° C to 185 ° C, most preferably about 178° C at the bottom of the column was used at the corresponding 1400 kPag column pressure.
For butene copolymers, at an appropriate height, between the overhead distillate and the entry location of the entrainer solvent, a side draw (not shown) can be withdrawn to remove the trans-2-butene isomer.
The side draw can be either a liquid or vapor and can be sent to downstream unit operations for further processing or to storage vessels.
The bottom stream 10 is concentrated with the entrainer solvent and the polymerization solvent. The composition may be from 80 to 95 M:\Trevor\TT Specs\9240-Cda(01-18).doc weight %, preferably 85 to 95 weight most preferably about 90% weight polymer solvent and 20 to 5 weight %, preferably 15 to 5 weight %, most preferably about 10% weight entrainer solvent. A portion of this stream, stream 11, is heated through a reboiler 105, and then returned to the bottom of the extraction column 103. The remainder, shown as stream 12, is fed to the third or down stream stripper column 107, In Figure 1, a 1o process heat exchanger 106, to recover the heat from stream 17 from the bottom of the stripper column 107 is shown. This heat integration configuration is optional and serves to illustrate the flexibility of the process to reduce energy operating costs.
Stream 13 exiting heat exchanger 106 contains approximately 90 weight % of polymer solvent and 10 weight % of entrainer solvent.
Separation of polymer solvent and entrainer solvent occurs by fractional distillation in column 107 with polymer solvent recovered in the overhead distillate stream 14 and the entrainer solvent at the bottom of stripper column 107 in stream 15. Stream 13 enters column 107 near the top, at approximately 1/5t" of the height of the column, measured from the top.
The feed location was found to be effective in the control of entrainer solvent that carries over into stream 14 and thereby reducing the purity of 3 o the polymer solvent.
The vapor overhead from column 107 is condensed in condenser unit 108, where reflux is returned to the column and the liquid distillate stream 14 is withdrawn. The reflux ratio was found to be an effective operating parameter to control the composition of the overhead distillate, stream 14, particularly the level of entrainer solvent. A reflux ratio (weight M:\Trevor\TT Specs19240-Cda(01-18).doc 11 ratio of reflux/liquid distillate) of 0.5:1 was found to be sufficient but may range from 0.1:1 to 10:1. A purity in excess of 99.9 weight % polymer solvent was obtained in stream 14. The quantity of stream 14 generated is sufficient for use in the catalyst preparation area of the polymerization process with any excess recycled directly to the polymerization reactors.
A purity of 99.6 weight % entrainer solvent in the bottom stream 15 1o was obtained. A portion of the bottom product stream 15 is heated through a reboiler, 109, and then returned to the bottom of the column, shown as stream 16. The remainder, shown as stream 17 is heat exchanged in the process/process exchanger 106, with the discharge directed to the top of the extraction column 103 to complete the recycle of the entrainer solvent stream. A small purge stream 18, is withdrawn to avoid accumulation of impurities in the entrainer solvent stream. Fresh makeup of entrainer solvent may be added back into the entrainer solvent recycle, shown as stream 8.
Any suitable operating pressure range in the stripper column 107 can be used to effect the separation. The pressure of 50 to 500 kPag, preferably from 130 kPag to 150 kPag , typically 140 to 150 (e.g. 145) kPag at the top of the column was used. The operating temperature 3o profile from 65° C to 90° C, typically from 68° C to 72° C, preferably about 70° C at the top and from 80° C to 110° C, preferably from 85° C to 90° C, typically about 87° C at the bottom of the column was used at the corresponding 145 kPag column top pressure.
For the production of hexene and octene copolymer, where 1-hexene and 1-octene are used respectively as the comonomer, the M:\Trevor\TT Specs\9240-Cda(01-18).doc 12 primary difference is where the polymerization solvent (PS) and the comonomer streams are recovered. In this scenario, a high purity polymerization solvent (PS) stream is recovered in the overhead of the extraction column 103, shown as stream 7. A high purity comonomer stream is recovered in the overhead of the stripper column 107, shown as stream 14. At an appropriate location below the stripper condenser, unit 108, a side draw can be taken to remove the comonomer isomers.
No additional distillation columns or changes in the process configuration are required to accommodate the use of 1-butene, 1-hexene, or 1-octene in the process of the present invention.
For the separation of hexene monomer from the polymerization solvent the extractive distillation column is typically operated at pressures from 50 to 500 kPag and a temperature profile at the top of the column from 70° C to 90° C, preferably from 75° C to 85°
C and a temperature profile at the bottom of the column from 170° C to 200° C, preferably from 180° C to 195° C. In the stripper column (107) the pressure is generally from 50 to 500 kPag and the temperature profile at the top of the column is from 65° C to 90° C, preferably from 70° C to 80°
C, and the temperature profile at the bottom of the column is from 170° C to 200° C, preferably from 180° C to 195° C.
For the separation of octene monomer from the polymerization solvent the extractive distillation column is typically operated at pressures from 50 to 500 kPag and a temperature profile at the top of the column from 65° C to 90° C, preferably from 75° C to 90°
C and a temperature profile at the bottom of the column from 130° C to 160° C, preferably from M:\Trevor\TT Specs\9240-Cda(01-18).doc 13 140° C to 155° C. In the stripper column (107) the pressure is generally from 50 to 500 kPag and the temperature profile at the top of the column is from 125° C to 145° C, preferably from 130° C to 140° C, and the temperature profile at the bottom of the column is from 200° C to 230° C, preferably from 210° C to 230° C.
The present invention will now be illustrated by the following non 1o limiting example. In the example unless otherwise stated the compositions are indicated in weight %.
A computer model developed using PRO/II, Simulation Sciences Inc. was used to model the process of the present invention. The model accurately reflects the solution polymerization processes of NOVA
Chemicals solution polymerization process within less than 3% error.
2o The extractive distillation process was modeled using a commercial process simulator to analyze the separation performance for a number of selected entrainers. In each case, a number of design and operating parameters were varied to determine the impact on product purity and recovery, including: feed tray location, reflux ratio, solvent circulation rate, and number of trays. These steady-state models were used to obtain a comparison of entrainer performance to arrive at a family/class of solvents 3o that show potential as an effective solvent for alkanelalkene separation and to develop a heat and material balance around the extraction process.
In order to develop a steady-state model of the extractive distillation process, thermodynamic methods need to be selected to characterize the vapor-liquid interaction between each component pairing in our system, M:\Trevor\TT Specs\9240-Cda(01-18).doc 14 particularly interaction with the entrainer component. In our study, the following two liquid activity coefficient methods were used:
~ Dortmund modification of UNIFAC
~ Non-Random Two Liquid (NRTL) The following references are suggested for a more detail treatment on the subject of thermodynamic characterization of excess functions using to liquid activity coefficient methods and heats of mixing.
~ The Dortmund modification of UNIFAC {J.Lohmann, et. al., "From UNIFAC to Modified UNIFAC (Dortmund)", Ind. Eng. Chem. Res., 2001, 40, pg. 957-964}.
~ NRTL method ~H.Renon, et. al., "Local Compositions in Thermodynamic Excess Functions For Liquid Mixtures", AIChE
Journal, 1, Vol. 14, 1968, pg. 135-144}.
The Dortmund modification of UNIFAC method provides a quick and reasonably accurate basis with which to perform a preliminary analysis of the effectiveness of various entrainers on product purity and recovery of alkane/alkene systems. Based on these results, the cyclic amide family/class of compounds (i.e. NMP) was selected as possible entrainers for further study.
To improve the level of accuracy in the extraction models, experimental measurement of the vapor-liquid equilibrium between the various binary pairings are needed, particularly for the solute pairings with the entrainer solvent. J. Gmehling, et al, has compiled an extensive collection of vapor-liquid equilibrium data that dates from 1977-present.
The DECHEMA series includes experimental data that covers a wide M:\Trevor\TT Specs\9240-Cda(01-18).doc 15 range of solute and solvent pairings, particularly those with commercial applications in the chemical industries. In addition, heats of mixing for the binaries are also reported along with activity coefficients at infinite dilution.
The latter is used to provide a metrics to rapidly compare and screen alternative entrainers for various separation systems.
Experimental data on vapor-liquid equilibrium of NMP with 1-to hexene and the components in our polymerization solvent has been published by K.Fischer and J.Gmehling (K.Fischer, et al., "Vapor-Liquid Equilibria, Activity Coefficients At Infinite Dilution and Heats of Mixing For Mixtures of N-Methyl Pyrrolidone-2 with C5 or C6 Hydrocarbons and for Hydrocarbon Mixtures", Fluid Phase Equilibria, 119, 1996, pg. 113-130.) that contains VLE data along with heats of mixing. The VLE data was 2o regressed to obtain coefficients for the 8-parameter form of the NRTL
equation for liquid activity coefficients for the various binary pairs. In each case, the regressed equation matches the experimental data over the entire composition range with an average deviation of 0.5%. Maximum deviation of 12.5% was observed at either ends of the curve (i.e. dilute regions).
Experimental heats of mixing data (K.Fischer, 1996) for various binary pairs was also regressed to the 8-parameter Redlich-Kister equation for excess heats of mixing. These parameters were used in the simulation model for the extraction column to provide a more accurate temperature profile across each tray. The maximum deviation between the regressed curve and the experimental data points is less than 1.6%.
M;\Trevor\TT Specs19240-Cda(01-18).doc 16 A computer model was developed using PRO/II, Simulation Sciences Inc., to model the extraction process using the NRTL parameters for k-value prediction and the Redlich-Kister parameters for the heats of mixing for liquid enthalpy prediction. The data is summarized in the three examples below.
EXAMPLES
Example 1 The feed system to be separated is a 50%wt 1-hexene/50% wt n-hexane mixture at a rate of 10,000 kg/hr using NMP as the entrainer solvent. The relative volatility between 1-hexene/n-hexane is merely 1.01 and is reflected in a difference of only 5° C in the normal boiling point. This mixture is a difficult one to separate by fractional distillation alone. The 2o model is used to provide an assessment of the selectivity and capacity of NMP along with the sensitivity on product purity and recovery with variations in entrainer:feed ratio, reflux ratio, and hexane product withdraw rate. Figure 2 shows the impact of entrainer:feed ratio on the purity of the hexane-rich stream and the hexene-rich stream. At an entrainer:feed weight ratio of 10:1, the purity of the hexane-rich stream recovered is 91 wt hexane with the remaining 9% wt being 1-hexene. The purity of the corresponding hexene-rich stream recovered is 84% wt 1-hexene with the remaining 16% wt being hexane. The purity of these two streams is more than sufficient for use in hexene solution copolymerization.
Figure 3 shows the impact of reflux ratio in the extraction column on the purity of the hexane-rich stream recovered, at three different product M:\TrevorlTT Specs\9240-Cda(01-18).doc 17 withdraw rates (i.e. hexane-rich stream at 4000, 4500, and 5000 kglh). In addition, Figure 3 shows the variation of NMP concentration in the hexane-rich stream at the various withdraw rates. As the hexane-rich withdraw rate increases, the concentration of NMP entrained decreases. The ppm levels detected are sufficiently low to enable recycle of the hexane-rich stream to the polymerization reaction.
to Figure 4 shows a 3-dimensional plot of the purity of the hexane-rich stream with variations in entrainer:feed ratio and hexane-rich withdraw rate.
The operating conditions for columns 100,103 and 107were substantially as shown in table 4 in example 3.
Example 2 For butene copolymer mode of operation, a butene/PS mixture is generated from polymerization. Using the extractive distillation process as described in this invention, a polymerization solvent-rich stream is generated (i.e. Stream 14) as overhead from column 107. In addition, a butene-rich stream is generated (i.e. Stream 7) as overhead from column 103. The purity of these two streams is summarized in Table 1. The operating conditions to achieve this separation are provided in Table 2.
M:lTrevor\TT Specs\9240-Cda(01-18).doc 1$
Table 1. Product Purity For Butene Mode Operation @ 0.1:1 Entrainer:Feed Ratio. Feed Rate=100 kglhr Composition Polymerization Butene Comonomer Solvent Recycle Recycle (Stream (Stream 14) 7) Ethylene (wt%) 0 0.8 Polymerization Solvent 99.99 3.3 (wt %) 1-Butene (wt%) 0 95
2-Butene (wt%) 0 1.0 NMP (ppm) 0 -. l O..
Table 2. Column Parameters For Butene Mode Operation @ 0.1:1 Entrainer:Feed Ratio, Feed Rate=100 kglhr Column 100 Column 103 Column 107 Theoretical Stages 10 35 20 Entrainer:Feed Ratio 0.1:1 (wt) Reflux Ratio (1Nt RefIux/Liquid1.5* 2 0.5 Distillate) Operating Pressure (Top,500-850 1200-1500 50-500 2o kPag) Operating Temp (Top, 90-115 90-110 65-85 C) Operating Temp (Bottom, 135-150 175-190 75-95 C) Notes:
Wt Reflux/Total Feed Example 3 For hexene copolymer mode of operation, a hexene/PS mixture is generated from polymerization. Using the extractive distillation process as described in this invention, a polymerization solvent-rich stream is generated (i.e. Stream 7) as overhead from column 103. In addition, a hexene-rich stream is generated (i.e. Stream 14) as overhead from column 107. The purity of these two streams is summarized in Table 3. The operating conditions to achieve this separation are provided in Table 4.
Without modifications to the configuration of the extractive distillation system, it can accommodate butene, hexene, or octene copolymerization.
M:\Trevor\TT Specs\9240-Cda(01-18).doc 1 g Table 3. Product Purity For Hexane Mode Operation @ 10:1 Entrainer:Feed Ratio, Feed Rate=100 kglhr.
Composition Polymerization Hexane Comonomer Solvent Recycle Recycle (Stream (Stream 7) 14) Ethylene (ppmw) 28 0 Polymerization Solvent 95 6.5 (wt %) 1-Hexane (wt%) 4.4 79 C6 Olefins (wt%) 0.3 14.5 NMP (ppb) -~ .<10 0 l0 Table 4. Column Parameters For Hexane Mode Operation @ 10:1 Entrainer:Feed Ratio, Feed Rate=100 kglhr.
Column 100 Column 103 Column 107 Theoretical Stages 10 35 20 Entrainer:Feed Ratio 10:1 (wt) Reflux Ratio (Vllt Reflux/Liquid1.2* 3 5 Distillate) Operating Pressure (Top,500-850 100-500 50-300 2o kPag) Operating Temp (Top, 125-135 70-85 65-85 C) Operating Temp (Bottom, 135-150 115-120 170-190 C) Notes:
~ Wt RefluxITotal Feed M:\Trevor\TT Specs\9240-Cda(01-18).doc 20
Table 2. Column Parameters For Butene Mode Operation @ 0.1:1 Entrainer:Feed Ratio, Feed Rate=100 kglhr Column 100 Column 103 Column 107 Theoretical Stages 10 35 20 Entrainer:Feed Ratio 0.1:1 (wt) Reflux Ratio (1Nt RefIux/Liquid1.5* 2 0.5 Distillate) Operating Pressure (Top,500-850 1200-1500 50-500 2o kPag) Operating Temp (Top, 90-115 90-110 65-85 C) Operating Temp (Bottom, 135-150 175-190 75-95 C) Notes:
Wt Reflux/Total Feed Example 3 For hexene copolymer mode of operation, a hexene/PS mixture is generated from polymerization. Using the extractive distillation process as described in this invention, a polymerization solvent-rich stream is generated (i.e. Stream 7) as overhead from column 103. In addition, a hexene-rich stream is generated (i.e. Stream 14) as overhead from column 107. The purity of these two streams is summarized in Table 3. The operating conditions to achieve this separation are provided in Table 4.
Without modifications to the configuration of the extractive distillation system, it can accommodate butene, hexene, or octene copolymerization.
M:\Trevor\TT Specs\9240-Cda(01-18).doc 1 g Table 3. Product Purity For Hexane Mode Operation @ 10:1 Entrainer:Feed Ratio, Feed Rate=100 kglhr.
Composition Polymerization Hexane Comonomer Solvent Recycle Recycle (Stream (Stream 7) 14) Ethylene (ppmw) 28 0 Polymerization Solvent 95 6.5 (wt %) 1-Hexane (wt%) 4.4 79 C6 Olefins (wt%) 0.3 14.5 NMP (ppb) -~ .<10 0 l0 Table 4. Column Parameters For Hexane Mode Operation @ 10:1 Entrainer:Feed Ratio, Feed Rate=100 kglhr.
Column 100 Column 103 Column 107 Theoretical Stages 10 35 20 Entrainer:Feed Ratio 10:1 (wt) Reflux Ratio (Vllt Reflux/Liquid1.2* 3 5 Distillate) Operating Pressure (Top,500-850 100-500 50-300 2o kPag) Operating Temp (Top, 125-135 70-85 65-85 C) Operating Temp (Bottom, 135-150 115-120 170-190 C) Notes:
~ Wt RefluxITotal Feed M:\Trevor\TT Specs\9240-Cda(01-18).doc 20
Claims (31)
1. A process for the extractive separation of a mixture of one or more C4-10 alkenes from a mixture of one or more C4-10 alkanes comprising feeding said mixture to a counter current extractive distillation column in which the extractive solvent comprises from 100 to 90 weight % of one or more polar C4-8 lactams and from 0 to 10 weight % of water and separating the mixture into at least one overhead low boiling stream and a higher boiling bottoms stream.
2. The process according to claim 1 wherein in said lactam is selected from the group consisting of N-methyl pyrrolidone, N-ethyl pyrrolidone, and N-propyl pyrrolidone.
3. The process according to claim 2, wherein said mixture is a middle stream from a first distillation column to which has been fed a mixture comprising at least one C4-10 alkene, at least one C2-3 alkene, one or more C1-10 alkane, and oligomers of said alkenes.
4. The process according to claim 3 wherein in feed to the first distillation column further comprises from 0 to 5 weight % of one or more members selected from the group consisting of hydrogen, carbon monoxide and carbon dioxide.
5. The process according to claim 3 wherein said first distillation column is a reflux column operated at a reflux ratio from 0.1:1 to 25:1.
6. The process according to claim 5, wherein in the first distillation column the pressure at the top of the column is from 50 to 1500 Kpag and the temperature at the top of said first distillation column is from 10°C to 100°C.
7. The process according to claim 6 wherein said C4-10 alkene is selected from the group consisting of 1-butene, 1- hexene and 1-octene.
8. The process according to claim 7 wherein said C1-10 alkane is selected from the group consisting of n-pentane, 2-methyl pentane, 3-methyl pentane, n-hexane, cyclohexane, 2-methyl hexane, 3-methyl hexane 2,2-dimethyl butane, 2,3-dimethyl butane, cyclopentane and mixtures thereof.
9. The process according to claim 8, wherein said lactam is N-methyl pyrrolidone.
10. The process according to claim 9, wherein the extractive distillation column has a condenser in the upper part of the column and operates with a reflux ratio between 0.1:1 and 20:1.
11. The process according to claim 10 wherein in the extractive distillation column the weight ratio of lactam to feed is from 0.01:1 to 20:1.
12. The method according to claim 11, wherein said C1-10 alkane comprises not less than 55 weight % of 2- methyl pentane.
13. The process according to claim 12, wherein the C4-8 alkene comprises not less than 80 weight % butene.
14 The process according to claim 13, wherein said C4-10 alkene is a mixture of 1-butene and 2-butene.
15. The process according to claim 9, wherein the pressure at the top of said extractive distillation column is from 1200 to 1600 Kpag, and the temperature profile in said extractive distillation column is from 75°C
to 100°C at the top of the column to 170°C to 190°C at the bottom of the column.
to 100°C at the top of the column to 170°C to 190°C at the bottom of the column.
16. The process according to claim 15, wherein two overheads are withdrawn from the extractive distillation column one consisting essentially of 1-butene and the other consisting essentially of 2-butene.
17. The process according to claim 16, wherein the bottom stream from said extractive distillation column comprises a mixture of from 85 to 95 weight % of alkane and from 5 to 15 weight % of said lactam.
18. The process according to claim 17, further comprising feeding said bottom stream to a third column to separate said lactam from said alkane.
19. The process according to claim 18, wherein said third column is operated at a pressure from 50 to 500 kPag at the top of the column and at temperature of from 65°C to 90°C at the top of the column and from 80°C to 110°C at the bottom of the column.
20. The process according to claim 19, wherein said third column is operated at a reflux ratio of 0.1:10 to 10:1.
21. The process according to claim 12, wherein said C4-10 alkene is selected from the group consisting of 1-hexene and 1-octene
22. The process according to claim 21 wherein from 85 to 99 weight %
of the overhead stream from the extractive distillation column is said C1-10 alkane.
of the overhead stream from the extractive distillation column is said C1-10 alkane.
23. The process according to claim 22, wherein the bottom stream from said extractive distillation column comprises a mixture of said alkene and said lactam.
24. The process according to claim 23, wherein said alkene consists not less than 80 weight % of hexene and the extractive distillation column is operated at a pressure of from 50 to 500 kPag and a temperature profile from 70°C to 90°C at the top of the column and from 170°C
to 200°C at the bottom of the column.
to 200°C at the bottom of the column.
25. The process according to claim 24, further including feeding the bottom stream from said extractive distillation column to a third column to separate said alkene from said lactam.
26 The process according to claim 25, wherein said third column is operated at a pressure from 50 to 500 kPag and at temperature profile of from 65°C to 90°C at the top of the column and from 170°C
to 200°C at the bottom of the column.
to 200°C at the bottom of the column.
27. The process according to claim 26, wherein said third column is operated at a reflux ratio of 0.1:10 to 10:1.
28. The process according to claim 23, wherein said alkene consists of not less than 80 weight % octene and the extractive distillation column is operated at a pressure of from 50 to 500 Kpag at the top of the column and a temperature profile from 65°C to 90°C at the top of the column and from 130°C to 160 °C at the bottom of the column.
29. The process according to claim 28, further including feeding the bottom stream from said extractive distillation column to a third column to separate said alkene from said lactam.
30 The process according to claim 29, wherein said third column is operated at a pressure from 50 to 500 kPag and at temperature profile of from 125°C to 145°C at the top of the column and from 200°C to 230°C at the bottom of the column.
31 The process according to claim 30, wherein said third column is operated at a reflux ratio of 0.1:10 to 10:1.
Priority Applications (1)
| Application Number | Priority Date | Filing Date | Title |
|---|---|---|---|
| CA 2387169 CA2387169A1 (en) | 2002-05-22 | 2002-05-22 | Extractive distillation |
Applications Claiming Priority (1)
| Application Number | Priority Date | Filing Date | Title |
|---|---|---|---|
| CA 2387169 CA2387169A1 (en) | 2002-05-22 | 2002-05-22 | Extractive distillation |
Publications (1)
| Publication Number | Publication Date |
|---|---|
| CA2387169A1 true CA2387169A1 (en) | 2003-11-22 |
Family
ID=29555366
Family Applications (1)
| Application Number | Title | Priority Date | Filing Date |
|---|---|---|---|
| CA 2387169 Abandoned CA2387169A1 (en) | 2002-05-22 | 2002-05-22 | Extractive distillation |
Country Status (1)
| Country | Link |
|---|---|
| CA (1) | CA2387169A1 (en) |
Cited By (3)
| Publication number | Priority date | Publication date | Assignee | Title |
|---|---|---|---|---|
| US20170320796A1 (en) * | 2014-10-31 | 2017-11-09 | Lg Chem, Ltd. | Distillation device |
| CN114736094A (en) * | 2021-01-07 | 2022-07-12 | 国家能源投资集团有限责任公司 | Method for separating alkane and olefin by liquid-liquid extraction |
| CN116199556A (en) * | 2021-11-30 | 2023-06-02 | 国家能源投资集团有限责任公司 | Method for separating mixture containing oxygen-containing compound, alkane and alkene |
-
2002
- 2002-05-22 CA CA 2387169 patent/CA2387169A1/en not_active Abandoned
Cited By (4)
| Publication number | Priority date | Publication date | Assignee | Title |
|---|---|---|---|---|
| US20170320796A1 (en) * | 2014-10-31 | 2017-11-09 | Lg Chem, Ltd. | Distillation device |
| US10464867B2 (en) * | 2014-10-31 | 2019-11-05 | Lg Chem, Ltd. | Distillation method |
| CN114736094A (en) * | 2021-01-07 | 2022-07-12 | 国家能源投资集团有限责任公司 | Method for separating alkane and olefin by liquid-liquid extraction |
| CN116199556A (en) * | 2021-11-30 | 2023-06-02 | 国家能源投资集团有限责任公司 | Method for separating mixture containing oxygen-containing compound, alkane and alkene |
Similar Documents
| Publication | Publication Date | Title |
|---|---|---|
| EP2454288B1 (en) | Polymerization process for olefin-based polymers | |
| US11002653B2 (en) | Process and apparatus to determine the degree of separation of two solution streams | |
| JPS6051130A (en) | Method of obtaining high purity butene-1 with small energy consumption | |
| Lei et al. | Process improvement on separating C4 by extractive distillation | |
| BR112018012874B1 (en) | PROCESS TO PRODUCE A POLYMERIC COMPOSITION | |
| KR20190055085A (en) | Process for efficient purging of polymer particles | |
| RU2734254C2 (en) | Methods and systems for producing 1,3-butadiene | |
| EP3720886A1 (en) | A method of recovering olefins in a solution polymerisation process | |
| US2917437A (en) | Fluid separation process control | |
| AU668771B2 (en) | Process for separating hydrocarbon mixtures by extractive distillation | |
| CA2387169A1 (en) | Extractive distillation | |
| EP3720887A1 (en) | A method of recovering olefins in a solution polymerisation process | |
| US4515661A (en) | Process for separating highly pure butene-1 and butene-2 | |
| US3368966A (en) | Control of isobutane fractionating column | |
| BR112020016056A2 (en) | PROCESS | |
| US3436437A (en) | Process for recovering high purity isoprene by extractive distillation with methoxypropionitrile | |
| US4010010A (en) | Gas-vapor separation process and apparatus with low energy consumption | |
| US3240833A (en) | Process and apparatus for treating hydrocarbon stream to remove contaminant | |
| EP3913006A1 (en) | Method for separating polybutene | |
| US3431315A (en) | Butene-1 purification | |
| JPS5826324B2 (en) | Encabinirno Seiseihou | |
| JPH0489868A (en) | Molecular weight fractionation wax from solid wax | |
| TW202104141A (en) | Process for removing polyunsaturated hydrocarbons from c4 hydrocarbon streams in the presence of mercaptans, disulfides and c5 hydrocarbons | |
| WO2007018516A1 (en) | Solvent-aided fractionation process | |
| TH29948B (en) | Polymerization process |
Legal Events
| Date | Code | Title | Description |
|---|---|---|---|
| EEER | Examination request | ||
| FZDE | Dead |