Note: Descriptions are shown in the official language in which they were submitted.
CA 02387169 2002-05-22
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/
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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.
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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
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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
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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
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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
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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.
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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.
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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
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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
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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
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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
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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,
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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
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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%.
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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
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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.
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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.
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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
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