U.S. patent application number 10/373587 was filed with the patent office on 2003-11-27 for continuous process for producing poly(trimethylene terephthalate).
Invention is credited to Giardino, Carl J., Griffith, David B., Ho, Chungfah Howard, Howell, James M., Watkins, Michelle Hoyt.
Application Number | 20030220465 10/373587 |
Document ID | / |
Family ID | 23998724 |
Filed Date | 2003-11-27 |
United States Patent
Application |
20030220465 |
Kind Code |
A1 |
Giardino, Carl J. ; et
al. |
November 27, 2003 |
Continuous process for producing poly(trimethylene
terephthalate)
Abstract
A continuous process for the production of poly(trimethylene
terephthalate) is disclosed. According to the process, a liquid
feed mixture comprising bis-3-hydroxypropyl terephthalate and/or
low molecular weight polyesters of 1,3-propanediol and terephthalic
acid, the liquid feed mixture-having a mole ratio of propylene
groups to terephthalate groups of 1.1 to 2.2 is fed to a flasher. A
first stream of gaseous by-products is continuously vaporized and
removed from the flasher, and a liquid flasher reaction product
having a mole ratio of propylene groups to terephthalate groups of
less than about 1.5 is continuously withdrawn from the flasher. The
liquid flasher reaction product is continuously fed to a
prepolymerizer where it is continuously polymerized to form a
poly(trimethylene terephthalate) prepolymer and a second stream of
gaseous by-products. Poly(trimethylene terephthalate) prepolymer
having a relative viscosity of at least about 5 is continuously
withdrawn from the prepolymerizer and continuously fed to a final
polymerizer, where it is continuously polymerized to form a higher
molecular weight poly(trimethylene terephthalate) and a third
stream of gaseous by-products. Higher molecular weight
poly(trimethylene terephthalate) having a relative viscosity of at
least about 17 is continuously withdrawn from the final
polymerizer.
Inventors: |
Giardino, Carl J.; (Hixson,
TN) ; Griffith, David B.; (Houston, TX) ; Ho,
Chungfah Howard; (Kinston, NC) ; Howell, James
M.; (Greenville, NC) ; Watkins, Michelle Hoyt;
(Waynesboro, VA) |
Correspondence
Address: |
E I DU PONT DE NEMOURS AND COMPANY
LEGAL PATENT RECORDS CENTER
BARLEY MILL PLAZA 25/1128
4417 LANCASTER PIKE
WILMINGTON
DE
19805
US
|
Family ID: |
23998724 |
Appl. No.: |
10/373587 |
Filed: |
February 25, 2003 |
Related U.S. Patent Documents
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Application
Number |
Filing Date |
Patent Number |
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10373587 |
Feb 25, 2003 |
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|
10057497 |
May 22, 2001 |
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6538076 |
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10057497 |
May 22, 2001 |
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09502642 |
Feb 11, 2000 |
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Current U.S.
Class: |
528/271 ;
528/272; 528/296 |
Current CPC
Class: |
C08G 63/78 20130101;
Y10T 428/2913 20150115; C08G 63/183 20130101; Y10T 428/31786
20150401 |
Class at
Publication: |
528/271 ;
528/272; 528/296 |
International
Class: |
C08G 063/00; C08G
069/00; C08G 067/00; C08G 063/02; C08G 063/12 |
Claims
1. A continuous process for the production of poly(trimethylene
terephthalate) comprising the steps of: (a) continuously feeding a
liquid feed mixture to a flasher, the liquid feed mixture
comprising a catalyst and at least one of bis-3-hydroxypropyl
terephthalate and low molecular weight polyesters of
1,3-propanediol and terephthalic acid, and the liquid feed mixture
having a mole ratio of propylene groups to terephthalate groups of
1.1 to 2.2; (b) continuously vaporizing and removing a first stream
of gaseous by-products from the flasher, and continuously
withdrawing a liquid flasher reaction product having a mole ratio
of propylene groups to terephthalate groups of less than about 1.5
from the flasher; (c) continuously feeding the liquid flasher
reaction product to a prepolymerizer, and continuously polymerizing
the flasher reaction product in the prepolymerizer to form a
poly(trimethylene terephthalate) prepolymer and a second stream of
gaseous by-products; (d) continuously withdrawing the
poly(trimethylene terephthalate) prepolymer from the
prepolymerizer, the prepolymer having a relative viscosity of at
least about 5; (e) continuously feeding the poly(trimethylene
terephthalate) prepolymer to a final polymerizer, and continuously
polymerizing the poly(trimethylene terephthalate) prepolymer to
form a higher molecular weight poly(trimethylene terephthalate) and
a third stream of gaseous by-products; and (f) continuously
withdrawing the higher molecular weight poly(trimethylene
terephthalate) from the final polymerizer, the higher molecular
weight poly(trimethylene terephthalate) having a relative viscosity
of at least about 17.
2. The process according to claim 1, wherein the temperature of
liquid reactants in the flasher is maintained at about 235.degree.
to about 250.degree. C.
3. The process according to claim 1, wherein the pressure in the
flasher is maintained at about 40 to about 80 mm of Hg.
4. The process according to claim 1, wherein the temperature of
liquid reactants in the prepolymerizer is maintained at about
240.degree. to about 255.degree. C.
5. The process according to claim 1, wherein the pressure in the
prepolymerizer is maintained at about 5 to about 30 mm of Hg.
6. The process according to claim 1, wherein the temperature of
liquid reactants in the final polymerizer is maintained at about
245.degree. to about 265.degree. C.
7. The process according to claim 1, wherein the pressure in the
final polymerizer is maintained at about 0.5 to about 3.0 mm
Hg.
8. The process according to claim 1, wherein the poly(trimethylene
terephthalate) that is withdrawn from the final polymerizer has a
relative viscosity of at least about 35.
9. The process according to claim 1, wherein the poly(trimethylene
terephthalate) that is withdrawn from the final polymerizer has a
relative viscosity of at least about 40.
10. The process according to claim 1, wherein the poly(trimethylene
terephthalate) that is withdrawn from the final polymerizer has a
relative viscosity of at least about 45.
11. The process according to claim 1, wherein the poly(trimethylene
terephthalate) that is withdrawn from the final polymerizer has a
relative viscosity of at least about 50.
12. The process according to claim 1, wherein the first stream of
gaseous by-products is continuously removed from the flasher and
condensed, the second stream of gaseous by-products is continuously
removed from the prepolymerizer and condensed, and the combined
first and second streams of condensed by-products contain not more
than 100 ppm by weight of acrolein and not more than 600 ppm of
allyl alcohol.
13. The process according to claim 12, wherein the combined first
and second streams of condensed by-products contain not more than
60 ppm by weight of acrolein and not more than 400 ppm of allyl
alcohol.
14. The process according to claim 12, wherein the combined first
and second streams of condensed by-products contain not more than
40 ppm by weight of acrolein and not more than 250 ppm of allyl
alcohol.
15. The process according to claim 1, wherein the third stream of
gaseous by-products is continuously removed from the final
polymerizer and condensed, and the third stream of condensed
by-products contains not more than 200 ppm of acrolein and not more
than 3000 ppm of allyl alcohol.
16. The process according to claim 15, wherein the third stream of
condensed by-products contains not more than 100 ppm of acrolein
and not more than 2500 ppm of allyl alcohol.
17. The process according to claim 15, wherein the third stream of
condensed by-products contains not more than 70 ppm of acrolein and
not more than 1000 ppm of allyl alcohol.
18. The process according to claim 1, wherein the temperature of
liquid reactants in the flasher is maintained at about 235.degree.
to about 250.degree. C.; the pressure in the flasher is maintained
at about 40 to about 80 mm of Hg; the temperature of liquid
reactants in the prepolymerizer is maintained at about 245.degree.
to about 255.degree. C.; the pressure in the prepolymerizer is
maintained at about 5 to about 30 mm of Hg; the temperature of
liquid reactants in the final polymerizer is maintained at about
245.degree. to about 265.degree. C.; and the pressure in the final
polymerizer is maintained at about 0.5 to about 3.0 mm Hg.
19. The process according to claim 1, wherein the temperature of
liquid reactants in the flasher is maintained at about 240.degree.
to about 245.degree. C.; the pressure in the flasher is maintained
at about 50 to about 70 mm of Hg; the temperature of liquid
reactants in the prepolymerizer is maintained at about 245.degree.
to about 250.degree. C.; the pressure in the prepolymerizer is
maintained at about 10 to about 20 mm of Hg; the temperature of
liquid reactants in the final polymerizer is maintained at about
250.degree. to about 260.degree. C.; and the pressure in the final
polymerizer is maintained at about 0.5 to about 3.0 mm Hg.
20. A continuous process for the production of poly(trimethylene
terephthalate) comprising the steps of: (a) continuously feeding a
liquid feed mixture to a flasher, the liquid feed mixture
comprising a catalyst and at least one of bis-3-hydroxypropyl
terephthalate and low molecular weight polyesters of
1,3-propanediol and terephthalic acid having an average degree of
polymerization of not greater than 15, and the liquid feed mixture
having a mole ratio of propylene groups to terephthalate groups of
1.1 to 2.2; (b) continuously vaporizing, removing and condensing a
first stream of gaseous by-products from the flasher; (c)
continuously withdrawing a liquid flasher reaction product having a
mole ratio of propylene groups to terephthalate groups of less than
about 1.5 from the flasher, wherein the temperature of liquid
reactants in the flasher is maintained at about 235.degree. to
about 250.degree. C., and the pressure in the flasher is maintained
at about 40 to about 80 mm of Hg; (d) continuously feeding the
liquid flasher reaction product to a prepolymerizer, and
continuously polymerizing the flasher reaction product in the
prepolymerizer to form a poly(trimethylene terephthalate)
prepolymer and a second stream of gaseous by-products, wherein the
temperature of liquid reactants in the prepolymerizer is maintained
at about 245.degree. to about 255.degree. C., and the pressure in
the prepolymerizer is maintained at about 5 to about 30 mm of Hg;
(e) continuously removing and condensing the second stream of
gaseous by-products from the prepolymerizer, whereby said first and
second streams of condensed by-products, when combined, contain not
more than 100 ppm by weight of acrolein and not more than 600 ppm
of allyl alcohol; (f) continuously withdrawing the
poly(trimethylene terephthalate) prepolymer from the
prepolymerizer, the prepolymer having a relative viscosity of at
least about 5; (g) continuously feeding the poly(trimethylene
terephthalate) prepolymer to a final polymerizer where the
poly(trimethylene terephthalate) prepolymer is continuously
polymerized to form a higher molecular weight poly(trimethylene
terephthalate) and a third stream of gaseous by-products, wherein
the temperature of liquid reactants in the final polymerizer is
maintained at 245.degree. to 265.degree. C., and the pressure in
the final polymerizer is maintained at 0.5 to 3.0 mm Hg; and (h)
continuously withdrawing the higher molecular weight
poly(trimethylene terephthalate) from the final polymerizer, the
higher molecular weight poly(trimethylene terephthalate) having a
relative viscosity of at least about 40; and (i) continuously
removing and condensing the third stream of gaseous by-products
from the final polymerizer, whereby the resulting condensate
contains not more than 200 ppm by weight of acrolein and not more
than 3000 ppm of allyl alcohol.
Description
FIELD OF THE INVENTION
[0001] The present invention relates to a continuous process for
the production of poly(trimethylene terephthalate), which is also
commonly referred to as poly(1,3-propylene terephthalate). The
process of the invention can be used as part of a four-vessel
process, the first vessel being either an ester exchanger for
producing a mixture of bis-3-hydroxypropyl terephthalate and low
molecular weight polymers of 1,3-propanediol and terephthalic acid
having an average degree of polymerization of 15 or less from
dimethylterephthalate and 1,3-propanediol or a reactor for
producing the starting material from terephthalic acid and
1,3-propanediol. The second vessel is a flasher, the third vessel
is a prepolymerizer, and the fourth vessel is a final polymerizer
or finisher.
BACKGROUND OF THE INVENTION
[0002] Continuous, four vessel processes for the production of
poly(ethylene terephthalate) are known. For example, Sheller, U.S.
Pat. No. 3,438,942 discloses a process for the continuous
production of poly(ethylene terephthalate) comprising ester
exchange followed by three polycondensation steps.
[0003] Also known are batch processes for the production of
poly(trimethylene terephthalate). For example, Doerr et al., U.S.
Pat. No. 5,340,909 discloses the production of poly(trimethylene
terephthalate) using either an ester exchange reaction starting
with lower dialkyl terephthalate ester or direct esterification of
terephthalic acid followed by a polycondensation reaction, both of
which are carried out in batches using an autoclave.
[0004] It would be highly desirable to provide a continuous,
four-vessel process for the production of poly(trimethylene
terephthalate). It would also be desirable to provide a continuous
process for the production of poly(trimethylene terephthalate) in
which the production of by-products, such as acrolein and allyl
alcohol, is minimized, and in which the molecular weight of the
final poly(trimethylene terephthalate) polymer is maximized. The
present invention provides such a process.
SUMMARY OF THE INVENTION
[0005] The present invention comprises a continuous process for the
production of poly(trimethylene terephthalate) comprising the steps
of:
[0006] (a) continuously feeding a liquid feed mixture to a flasher,
the liquid feed mixture comprising a catalyst and at least one of
bis-3-hydroxypropyl terephthalate and low molecular weight
polyesters of 1,3-propanediol and terephthalic acid, and the liquid
feed mixture having a mole ratio of propylene groups to
terephthalate groups of 1.1 to 2.2;
[0007] (b) continuously vaporizing and removing a first stream of
gaseous by-products from the flasher, and continuously withdrawing
a liquid flasher reaction product having a mole ratio of propylene
groups to terephthalate groups of less than about 1.5 from the
flasher;
[0008] (c) continuously feeding the liquid flasher reaction product
to a prepolymerizer, and continuously polymerizing the flasher
reaction product in the prepolymerizer to form a poly(trimethylene
terephthalate) prepolymer and a second stream of gaseous
by-products;
[0009] (d) continuously withdrawing the poly(trimethylene
terephthalate) prepolymer from the prepolymerizer, the prepolymer
having a relative viscosity of at least about 5;
[0010] (e) continuously feeding the poly(trimethylene
terephthalate) prepolymer to a final polymerizer, and continuously
polymerizing the poly(trimethylene terephthalate) prepolymer to
form a higher molecular weight poly(trimethylene terephthalate) and
a third stream of gaseous by-products; and
[0011] (f) continuously withdrawing the higher molecular weight
poly(trimethylene terephthalate) from the final polymerizer, the
higher molecular weight poly(trimethylene terephthalate) having a
relative viscosity of at least about 17.
DESCRIPTION OF THE DRAWINGS
[0012] FIG. 1 is a schematic representation of an apparatus useful
in carrying out the process of the invention.
DETAILED DESCRIPTION OF THE INVENTION
[0013] The process of the invention is part of a continuous,
four-vessel, four-stage process for the production of
poly(trimethylene terephthalate). The first stage in the process is
either an ester exchange or direct esterification reaction,
depending upon whether the starting material for the process is
dimethylterephthalate or terephthalic acid. The second stage is the
rapid removal of 1,3-propanediol in a flasher, the third stage is a
prepolymerization, and the fourth stage is a final polymerization.
The present invention is useful for the production of
poly(trimethylene terephthalate) containing low levels of toxic
byproducts such as acrolein and allyl alcohol.
[0014] The term "ppm" is used herein to mean parts per million and
is equal to micrograms per gram.
[0015] 1. Production of Feed Materials
[0016] The feed material for the flasher may be produced either by
ester exchange from dimethylterephthalate and 1,3-propanediol or by
direct esterification from terephthalic acid and 1,3-propanediol.
Both processes yield bis-3-hydroxypropyl terephthalate (referred to
as "monomer") and low molecular weight polyesters of
1,3-propanediol and terephthalic acid having an average degree of
polymerization of 15 or less (referred to as "oligomers").
[0017] As shown in FIG. 1, reaction vessel 10 is a source of
monomer and/or oligomers, which are fed to flasher 12. Reaction
vessel 10 can be either an ester exchange reactor or a direct
esterification reactor.
[0018] Whether the monomer/oligomer feed mixture is produced by
direct esterification from terephthalic acid or ester exchange from
dimethylterephthalate, a catalyst is added prior to the
esterification or transesterification reaction. Catalysts useful in
the ester exchange process include organic and inorganic compounds
of titanium, lanthanum, and zinc. Titanium catalysts, such as
tetraisopropyl titanate and tetraisobutyl titanate are preferred
and are added to the 1,3-propanediol in an amount sufficient to
yield 20 to 90 ppm of titanium by weight based on the finished
polymer. These levels produce relatively low unreacted
dimethylterephthalate in the ester exchange reaction (less than 5%
by weight based on the total weight of the exit stream from the
ester exchange), give reasonable reaction rates in the
prepolymerization and final polymerization steps, and produce
polymer with CIELAB b* color of less than 8 as measured by the CIE
1976 CIELAB color scale as standardized by CIE, the Commission
International de L'Eclairage. The b-value shows the degree of
yellowness, with a higher value showing a higher (undesirable)
degree of yellowness. Another useful ester exchange catalyst is
lanthanum acetate, which may be added in an amount sufficient to
yield 125 to 250 ppm of lanthanum by weight based on the finished
polymer. Following the ester exchange reaction, the lanthanum is
deactivated by the addition of phosphoric acid in an amount
sufficient to yield 10 to 50 ppm of phosphorus by weight based on
the finished polymer. Tetraisopropyl titanate or tetraisobutyl
titanate is then added as a polycondensation catalyst in an amount
sufficient to yield 10 to 50 ppm of titanium by weight based on the
finished polymer. Amounts of other ester exchange catalysts are
adjusted to give the same effect as the 20 to 90 ppm of
titanium.
[0019] Catalysts useful in the direct esterification process
include organo-titanium and organo-tin compounds, which are added
to the 1,3-propanediol in an amount sufficient to yield at least 20
ppm of titanium, or at least 50 ppm of tin, respectively, by weight
based on the finished polymer.
[0020] Additional catalyst may be added to the monomer/oligomer
mixture after the ester exchange or direct esterification reaction
and prior to prepolymerization.
[0021] Whether the monomer/oligomer feed mixture is produced by
direct esterification from terephthalic acid or ester exchange from
dimethylterephthalate, the mole ratio of propylene groups to
terephthalate groups is maintained at about 1.1 to 2.2, preferably
about 1.4 to 1.8, and more preferably about 1.5 entering the
flasher.
[0022] 2. Flasher As shown in FIG. 1, the monomer/oligomer mixture
is pumped from the ester exchanger or direct esterification reactor
to flasher 12 by means of a temperature-controlled feed line 11
equipped with pumps and filters. In the feed lines, the
monomer/oligomer mixture is maintained at a temperature of about
215.degree. to 250.degree. C.
[0023] The flasher is a jacketed and heated vessel with an internal
heater. The internal heater heats and vaporizes the excess
1,3-propanediol in the feed material. The bubbling of the
1,3-propanediol vapor provides the needed agitation. The excess
1,3-propanediol is removed through vapor line 13 connected to a
vacuum source and then condensed. In the flasher, the
monomer/oligomer mixture is maintained at a temperature of about
235.degree. to 250.degree. C., preferably about 240.degree. to
245.degree. C., and more preferably about 245.degree. C. The
pressure in the flasher is maintained at about 40 to 80 mm of Hg
(5332 to 10,664 Pa), preferably about 45 to 75 mm Hg (5998 to 9998
Pa), and more preferably about 50 to 70 mm of Hg (6665 to 9331
Pa).
[0024] In the flasher, the monomer/oligomer mixture reacts to form
a liquid flasher reaction product comprising a low molecular weight
trimethylene terephthalate polymer and releasing 1,3-propanediol as
a by-product. The excess 1,3-propanediol is vaporized and
continuously removed from the liquid reactants, lowering the
1,3-propanediol to dimethylterephthalate mole ratio to less than
about 1.5, preferably less than about 1.3, in the liquid flasher
reaction product.
[0025] The excess 1,3-propanediol that is removed from the flasher
can be condensed by means of spray condenser 14. Vapors from vapor
line 13 pass into a vertical condenser, where they are sprayed with
condensed 1,3-propanediol that has been cooled to a temperature of
less than 60.degree. C., preferably less than 50.degree. C. The
condensed 1,3-propanediol vapors from flasher 12, together with the
1,3-propanediol spray, flow into hotwell 15 located beneath
condenser 14, where they are combined with additional
1,3-propanediol. A portion of the liquid mixture in hotwell 14 is
pumped through a cooler to the top of the condenser for use as the
condensing spray. The condensed vapors from flasher 12 are combined
with the condensed vapors from prepolymerizer 17 in hotwell 15.
[0026] 3. Prepolymerization
[0027] As shown in FIG. 1, the flasher reaction product is fed via
temperature-controlled feed line 16 to prepolymerizer 17.
Prepolymerizer 17 performs the initial polymerization step, which
involves removing excess 1,3-propanediol and increasing the product
viscosity by building longer chain molecules of polymer.
[0028] The prepolymerizer is a jacketed and heated vessel with an
internal agitator. The agitator provides agitation and creates
liquid/vapor surface area for 1,3-propanediol removal. The
temperature of liquid reactants in the prepolymerizer is maintained
at about 240.degree. to 255.degree. C., preferably about
245.degree. to. 250.degree. C., and more preferably about
250.degree. C. The pressure in the prepolymerizer is maintained at
about 5 to 30 mm of Hg (666 to 3999 Pa), preferably about 10 to 20
mm of Hg (1333 to 2666 Pa), and more preferably about 15 mm of Hg
(1999 Pa).
[0029] The excess 1,3-propanediol is removed through vapor line 18
connected to a vacuum source and then condensed. One method for
condensing the 1,3-propanediol vapors from the prepolymerizer is by
means of spray condenser 19 similar to that described above for
condensing 1,3-propanediol vapors from the flasher. The condensed
vapors from prepolymerizer 17 are combined with the condensed
vapors from flasher 12 in hotwell 15.
[0030] The condensed 1,3-propanediol vapors exiting the flasher and
prepolymerizer typically contain other reaction by-products such as
acrolein and allyl alcohol. It is desirable that the production of
by-products such as acrolein and allyl alcohol be minimized because
both of these compounds are highly toxic and cause irritation to
the eyes and mucous membranes. According to the process of the
invention, the amount of acrolein contained in the combined
condensed 1,3-propanediol streams exiting the flasher and
prepolymerizer is no greater than 100 ppm by weight of condensate,
preferably no greater than 60 ppm, and more preferably no greater
than 40 ppm. The amount of allyl alcohol contained in the combined
condensed 1,3-propanediol streams exiting the flasher and
prepolymerizer is no greater than 600 ppm by weight of condensate,
preferably no greater than 400 ppm, and more preferably no greater
than 250 ppm.
[0031] Relative viscosity is an indicator of molecular weight.
Relative viscosity, often referred to as "LRV," is the ratio of the
viscosity of a solution of 4.75 grams of poly(trimethylene
terephthalate) in 100 grams of solution to the viscosity of the
solvent itself. The solvent used herein for measuring relative
viscosity is hexafluoroisopropanol containing 100 ppm sulfuric
acid, and the measurements are made at 25.degree. C. The
poly(trimethylene terephthalate) prepolymer that is withdrawn from
the prepolymerizer has a relative viscosity of at least about 5,
preferably about 5.5 to 7.
[0032] The residence or hold-up time in the prepolymerizer
typically ranges from about 30 to 90 minutes.
[0033] 4. Final Polymerization
[0034] As shown in FIG. 1, the liquid reaction product from
prepolymerizer 17 is fed via temperature-controlled feed line 20 to
final polymerizer or finisher 21. The major purpose of finisher 21
is to increase the molecular chain length or viscosity of the
polymer. This is accomplished by using heat, agitation, vacuum and
catalyst. It is desirable that the molecular weight of the finished
polymer be maximized, so that further processing, e.g., solid state
polymerization, can be avoided prior to fiber spinning or other
forming operation.
[0035] The finisher is normally a horizontal cylindrical vessel
surrounded by a jacket containing a heating medium, such as
Dowtherm vapor. Prepolymer from prepolymerizer 17 flows through an
inlet into the finisher. An agitator generates large surface areas
of thin films of polymer to enhance the mass transfer of
1,3-propanediol from the polymer.
[0036] The temperature of the liquid reactants in the finisher is
maintained at about 245.degree. to 265.degree. C., preferably about
250.degree. to 260.degree. C., and more preferably about
255.degree. C. The pressure in the finisher is maintained at about
0.5 to 3.0 mm Hg (66 to 399 Pa).
[0037] Finished polymer is removed from the finisher through an
outlet by means of a pump. The relative viscosity of the
poly(trimethylene terephthalate) exiting the finisher is at least
about 17, preferably at least about 35, more preferably at least
about 40, more preferably at least about 45, and most preferably at
least about 50. When correlated to intrinsic viscosity measurements
in 60/40 weight percent phenol/1,1,2,2-tetrachloroethane following
ASTM D 4603-96, these relative viscosities correspond to intrinsic
viscosities of 0.55 dl/g, 0.85 dl/g, 0.91 dl/g, 0.96 dl/g, and 1.0
dl/g, respectively. The viscosity of the finished polymer may be
controlled by adjusting finisher pressure or other variables. The
residence or hold-up time in the finisher is typically about 1 to 2
hours.
[0038] 1,3-Propanediol and other gaseous by-products are removed
from the finisher through vapor line 22 connected to a vacuum
source and then condensed. One method for condensing the
1.3-propanediol vapors from the finisher is by means of spray
condenser 23 similar to that described above for condensing
1,3-propanediol vapors from the flasher and prepolymerizer. The
condensed vapors from finisher 21 are collected in hotwell 24.
[0039] According to the present invention, the amount of acrolein
contained in the condensed 1,3-propanediol stream exiting the
finisher is no greater than 200 ppm by weight of condensate,
preferably no greater than 100 ppm, and more preferably no greater
than 70 ppm. The amount of allyl alcohol contained in the condensed
1,3-propanediol stream exiting the finisher is no greater than 3000
ppm, preferably no greater than 2500 ppm, and more preferably no
greater than 1000 ppm.
[0040] The finished polymer may be pelletized or fed directly to a
forming operation, such as fiber spinning, film formation or
molding operation. Fibers made from the poly(trimethylene
terephthalate) produced by the process of the invention have
properties which make them useful in various textile applications,
including the manufacture of carpet or apparel.
[0041] 5. Additives
[0042] Various additives may be used in the process of the
invention. These include color inhibitors, such as phosphoric acid,
delustrants, such as titanium dioxide, dyeability modifiers,
pigments and whiteners. If separate ester exchange and
polymerization catalysts are used, phosphoric acid
(H.sub.3PO.sub.4) or other color inhibitors may be added to
minimize or prevent the color forming property of the ester
exchange catalyst.
EXAMPLES 1 to 10
[0043] Poly(trimethylene terephthalate) was prepared using an
apparatus of the type indicated in the drawing, including an ester
exchanger, a flasher, a prepolymerizer and a finisher. In Examples
1-8, a 94.1 lb./hr (42.7 kg/hr) stream of dimethylterephthalate was
preheated to a temperature of 185.degree. C. and continuously mixed
with a 55.3 lb./hr (25.1 kg/hr) stream of catalyzed 1,3-propanediol
which was also preheated to a temperature of 185.degree. C., to
form a mixture having a mole ratio of 1.5 moles of 1,3-propanediol
per mole of dimethylterephthalate. In Example 9, the throughput was
lowered to 51.4 lb./hr (23.3 kg/hr) of dimethylterephthalate and
40.3 lb./hr (18.3 kg/hr) of catalyzed 1,3-propanediol which were
combined to form a mixture having a mole ratio of 2.0 moles of
1,3-propanediol per mole of dimethylterephthalate. In Example 10,
the throughput was lowered still further to 38.2 lb./hr (17.3
kg/hr) of dimethylterephthalate and 30.0 lb./hr (13.6 kg/hr) of
catalyzed 1,3-propanediol which were combined to form a mixture
having a mole ratio of 2.0 moles of 1,3-propanediol per mole of
dimethylterephthalate. The catalyst was tetraisopropyl titanate
(Tyzor.RTM. TPT, available from E. I. du Pont de Nemours and
Company, Wilmington, Del.). In Examples 1-8, the tetraisopropyl
titanate was added to the 1,3-propanediol in an amount sufficient
to yield 30-60 ppm by weight of titanium based on the weight of
poly(trimethylene terephthalate) formed in the process. In Examples
9 and 10, the catalyst level was raised to 70 ppm of titanium. The
dimethylterephthalate/catalyzed 1,3-propanediol mixture was fed
into the base of an ester exchanger, where the pressure at the base
of the ester exchanger was maintained at 825 to 900 mm of Hg
(109,972 to 119,970 Pa). In Examples 1-8, the temperature of the
liquid reactants in the ester exchanger was maintained at
230.degree. C., and in Examples 9 and 10, the temperature of liquid
reactants in the ester exchanger was maintained at 237.degree. C.
and 239.degree. C., respectively. The pressure at the top of the
ester exchange column was atmospheric. In the ester exchanger, the
1,3-propanediol reacted with the dimethylterephthalate to form
bis-3-hydroxypropyl terephthalate monomer and low molecular weight
oligomers of 1,3-propanediol and terephthalic acid, liberating
methanol vapor, which was continuously removed from the top of the
ester exchanger. The monomer/oligomer mixture was continuously
removed from the base of the ester exchanger and fed to the inlet
of a flasher. In the flasher, the monomers and oligomers reacted to
form a low molecular weight trimethylene terephthalate polymer,
liberating 1,3-propanediol vapor. The 1,3-propanediol vapor and
other gaseous by-products were removed from the top of the flasher
and condensed. The low molecular weight trimethylene terephthalate
polymer was continuously withdrawn from the flasher and fed to the
inlet end of a prepolymerizer. In the prepolymerizer, the monomers
and oligomers further reacted to form a higher molecular weight
poly(trimethylene terephthalate) prepolymer, liberating
1,3-propanediol vapor. The 1,3-propanediol vapor and other gaseous
by-products were removed from the top of the prepolymerizer,
condensed and combined with the condensates from the flasher. The
poly(trimethylene terephthalate) prepolymer was continuously
withdrawn from the prepolymerizer and fed to the inlet end of a
finisher vessel. The temperature of the liquid reactants in the
finisher was maintained at 255.degree. to 260.degree. C. In the
finisher, the poly(trimethylene terephthalate) prepolymer reacted
to form an even higher molecular weight polymer, liberating
additional 1,3-propanediol vapor. The 1,3-propanediol vapor and
other gaseous by-products were continuously removed from the
finisher. The poly(trimethylene terephthalate) was continuously
removed from the finisher and pelletized. The conditions and
results for the continuous polymerization are set forth in Tables
I, II and III. In Examples 9 and 10, the levels of polymer and
hold-up times in the finisher were reduced, resulting in lower
by-product formation and higher relative viscosity (LRV).
[0044] In the Tables, the acrolein and allyl alcohol levels are
given in parts per million (ppm) by weight based on the combined
condensates that are removed from the flasher and prepolymerizer
and the condensates that are removed from the finisher,
respectively. The dipropylene glycol (DPG) levels are given as a
weight percent based on the total prepolymer or finished polymer
that is removed from the flasher, prepolymerizer and finisher,
respectively. The speed of the agitator in the finisher is given in
revolutions per minute (RPM). The amount of carboxyl end groups
(COOH) in the finished polymer is given in microequivalents per
gram based on the total weight of the finished polymer. The level
of catalyst is given as parts per million (ppm) by weight of
titanium in the finished polymer.
1 TABLE I CATALYST FLASHER Ti Temperature Pressure 3G/T COOH DPG
EXAMPLE (ppm) (.degree. C.) mm Hg (Pa) mole ratio Microeq./g (wt.
%) 1 50 245 60 (7998) 1.22 1.9 0.18 2 40 245 60 (7998) 1.29 1.8
0.16 3 50 245 60 (7998) 1.08 1.4 0.15 4 60 245 60 (7998) 1.24 1.4
0.14 5 50 245 60 (7998) 1.18 1.4 0.13 6 30 245 60 (7998) 1.09 2.9
0.14 7 30 245 60 (7998) 1.19 1.6 0.14 8 30 245 60 (7998) 1.17 1.3
0.13 9 70 245 50 (6665) 1.51 2.6 10 70 245 50 (6665) 1.42 5.6
[0045]
2 TABLE II FLASHER/ PREPOLYMERIZER PREPOLYMERIZER Allyl Temp.
Pressure DPG COOH Acrolein Alcohol EXAMPLE (.degree. C.) mm Hg (Pa)
LRV (wt. %) microeq./g (ppm) (ppm) 1 250 15 (1999) 6.7 0.19 2.3 15
410 2 250 15 (1999) 6.6 0.16 2.4 107 516 3 250 15 (1999) 6.7 0.16
2.0 62 453 4 250 15 (1999) 5.9 0.15 2.2 69 526 5 250 30 (3999) 5.5
0.14 1.6 39 544 6 250 39 (5199) 5.0 0.15 1.8 76 565 7 250 20 (2666)
5.9 0.14 1.7 56 568 8 250 40 (5332) 5.4 0.13 1.5 90 525 9 250 15
(1999) 5.7 3.4 66 294 10 250 15 (1999) 5.9 3.1 63 299
[0046]
3 TABLE III FINISHER Agitator COOH Allyl Temp. Pressure Speed DPG
(micro- Acrolein Alcohol EXAMPLE (.degree. C.) mm Hg (Pa) (rpm) LRV
(wt. %) eq./g) (ppm) (ppm) 1 255 <5 (<666) 3 35 0.20 19 136
2848 2 255 <5 (<666) 3 35 0.23 20 77 2890 3 255 <5
(<666) 3.6 35 0.20 19 129 2778 4 255 <5 (<666) 3.6 35 0.19
22 0 2400 5 255 <5 (<666) 4 31 0.17 12 85 2569 6 255 <5
(<666) 4 31 0.18 12 0 2551 7 260 <5 (<666) 4 30 0.17 15 93
2674 8 260 <5 (<666) 4 32 0.17 18 0 3093 9 255 1.4 (187) 2 46
11 26 413 10 255 1.4 (187) 2 52 12 25 427
* * * * *