U.S. patent application number 13/635954 was filed with the patent office on 2013-08-08 for process for the production of water and solvent-free halobutyl rubbers.
This patent application is currently assigned to LANXESS INTERNATIONAL SA. The applicant listed for this patent is Rolf Feller, Adam Gronowski, Jorg Kirchhoff, John Lovegrove, Hanns-Ingolf Paul, Udo Wiesner. Invention is credited to Rolf Feller, Adam Gronowski, Jorg Kirchhoff, John Lovegrove, Hanns-Ingolf Paul, Udo Wiesner.
Application Number | 20130203943 13/635954 |
Document ID | / |
Family ID | 44140897 |
Filed Date | 2013-08-08 |
United States Patent
Application |
20130203943 |
Kind Code |
A1 |
Kirchhoff; Jorg ; et
al. |
August 8, 2013 |
PROCESS FOR THE PRODUCTION OF WATER AND SOLVENT-FREE HALOBUTYL
RUBBERS
Abstract
The present invention relates to water and solvent-free
halogenated butyl rubber products as a process for the production
thereof. The process comprises at least the steps of: a.) treating
a fluid containing at least one halogenated butyl rubber and at
least one volatile compound in at least one concentrator unit, b.)
reheating the concentrated fluid and c.) feeding the reheated
concentrated fluid into at least one extruder.
Inventors: |
Kirchhoff; Jorg; (Koln,
DE) ; Gronowski; Adam; (Sarnia, CA) ;
Lovegrove; John; (Sarnia, CA) ; Paul;
Hanns-Ingolf; (Leverkusen, DE) ; Feller; Rolf;
(Mettmann, DE) ; Wiesner; Udo; (Bornheim,
DE) |
|
Applicant: |
Name |
City |
State |
Country |
Type |
Kirchhoff; Jorg
Gronowski; Adam
Lovegrove; John
Paul; Hanns-Ingolf
Feller; Rolf
Wiesner; Udo |
Koln
Sarnia
Sarnia
Leverkusen
Mettmann
Bornheim |
|
DE
CA
CA
DE
DE
DE |
|
|
Assignee: |
LANXESS INTERNATIONAL SA
Granges-Paccot
CH
|
Family ID: |
44140897 |
Appl. No.: |
13/635954 |
Filed: |
March 23, 2011 |
PCT Filed: |
March 23, 2011 |
PCT NO: |
PCT/EP11/54447 |
371 Date: |
April 26, 2013 |
Current U.S.
Class: |
525/332.3 |
Current CPC
Class: |
B29C 48/767 20190201;
C08C 19/12 20130101; C08F 6/003 20130101; C08F 6/008 20130101; B29B
7/007 20130101; B29B 7/603 20130101; B29B 7/86 20130101; B29B 7/845
20130101; B29B 7/7495 20130101; C08F 6/10 20130101; B29C 48/76
20190201; C08F 6/003 20130101; C08L 23/283 20130101 |
Class at
Publication: |
525/332.3 |
International
Class: |
C08F 6/10 20060101
C08F006/10 |
Foreign Application Data
Date |
Code |
Application Number |
Mar 24, 2010 |
EP |
10003140.0 |
Mar 25, 2010 |
EP |
10157706.2 |
Claims
1. Process of removing volatile compounds from a fluid (F)
containing at least one halogenated butyl rubber and at least one
volatile compound which comprises at least the steps of: a)
treating the fluid (F) in at least one concentrator unit comprising
at least a heater (2), a degassing vessel (4) and a vapor line
(4.1), whereby the fluid (F) is heated, the heated fluid (G) is fed
into a degassing vessel (4) where part of the volatile compounds
are removed via the vapor line (4.1) to obtain a concentrated fluid
(H), b) reheating the concentrated fluid (H) from step a) in at
least one reheating unit (6) to obtain a reheated concentrated
fluid (L); c) feeding the reheated concentrated fluid (L) from step
b) into at least one extruder unit comprising at least an extruder
degassing section comprising at least a conveying section (16), a
vent port (15) with one or more vapor lines (15.1), a accumulating
section (20) and an outlet section (22), whereby volatile compounds
are removed through the vent ports (15) and vapor lines (15.1);
whereby the product (P) obtained at the outlet section (22) is
substantially five of volatile compounds.
2. The process pursuant to claim 1, characterized in that the
reheated concentrated fluid (L) is free-flowing.
3. The process pursuant to claim 1, characterized in that the
halogenated butyl rubber is bromobutyl rubber.
4. The process pursuant to claims 1 to 3, characterized in that the
Fluid (F) contains from 3 to 50 wt % of a halogenated butyl rubber
and from 60 to 97 wt % volatile compounds, whereby the
aforementioned components add up to 90 to 100 of the total mass of
fluid (F).
5. The process pursuant to claims 1 to 4, characterized in that the
degassing vessel (4) is designed in the shape of a cyclone and has
at least a torisperical shaped bottom to facilitate removal of
concentrated fluid (H).
6. The process pursuant to claims 1 to 5, characterized in that the
extruder unit comprises an extruder selected from the group
consisting of single and multiscrew extruders.
7. The process pursuant to claims 1 to 6, characterized in that the
extruder unit comprises means to operate separate zones of the
extruder independently of each other at different temperatures so
that the zones can either be heated, unheated or cooled.
8. The process pursuant to claims 1 to 7, characterized in that the
vent ports (15) comprise means to prevent the reheated concentrated
fluid (L) or the Product (P), from coming out of the vent ports and
that the means are stuffer screws.
9. The process pursuant to claims 1 to 8, characterized in that a
stripping agent is added in the extruder unit.
10. The process pursuant to claims 1 to 9, characterized in that
the extruder unit comprises at least one extruder degassing section
in upstream direction.
11. The process pursuant to claims 1 to 10, characterized in that
fluid (F) is obtained by a process of removing hydrophilic
compounds and optionally water from a crude fluid (A) containing at
least one non-volatile polymer, at least one volatile organic
compound, one or more hydrophilic compounds and optionally water
which comprises at least the step of pre a) treating the crude
fluid (A) in at least one pre-washing unit comprising at least a
separating apparatus (26), whereby the fluid (A) is mixed with
water to obtain an organic phase (28) comprising primarily
non-volatile polymer and volatile organic compounds and an aqueous
phase (27) comprising primarily water and hydrophilic compounds,
and whereby the organic phase (28) is separated from the aqueous
phase (27) in a separating apparatus (26) and further used as fluid
F and whereby at least a part of the aqueous phase (27) is removed
from the separating apparatus (fluid C).
12. The process pursuant to claims 1 to 11, characterized in that
fluid (A) or fluid (F) is obtained by a process comprising at least
the steps of I) providing a reaction medium comprising a common
aliphatic medium comprising at least 50 wt.-% of one or more
aliphatic hydrocarbons having a boiling point in the range of
45.degree. C. to 80.degree. C. at a pressure of 1013 hPa, and a
monomer mixture comprising at least one isoolefin monomer, at least
one multiolefin monomer and either no or at least one other
co-polymerizable monomer in a mass ratio of monomer mixture to
common aliphatic medium of from 40:60 to 99:1 II) polymerizing the
monomer mixture within the reaction medium to form a rubber
solution comprising a rubber polymer which is at least
substantially dissolved in the medium comprising the common
aliphatic medium and residual monomers of the monomer mixture; III)
separating residual monomers of the monomer mixture from the rubber
solution to form a separated rubber solution comprising the rubber
and the common aliphatic medium, IV) halogenating the rubber in the
separated rubber solution using a halogenating agent.
13. The process pursuant to claim 12, characterized in that the
halogenating agent is a brominating agent.
14. The process pursuant to claim 13, characterized in that the
brominating agent is at least partially regenerated by an oxidizing
agent.
Description
[0001] The present invention relates to water and solvent-free
halogenated butyl rubber products as well as a process for the
production thereof.
[0002] Halogenated butyl rubbers have important industrial uses and
are typically produced by the (co)polymerization of monomers, which
is typically carried out via slurry, emulsion or solution
processes.
[0003] Copolymerization of isobutene and isoprene, which leads to
butyl rubber, for example is carried out industrially at low
temperatures of approximately -60.degree. C. to -100.degree. C. to
obtain high molar masses. The slurry process uses chloromethane as
a diluent while the solution process uses an inert hydrocarbon as a
solvent. After the polymerization, the butyl rubber polymer is
present either as a slurry in chloromethane or as a homogeneous
solution in a hydrocarbon. Unreacted monomers are also present in
the reactor discharge mixture. The butyl rubber polymer needs to be
recovered and isolated from the diluent or solvent.
[0004] After the (co)polymerization, the reactor discharge mixture
contains at least the butyl rubber, solvents, residual monomers and
the catalyst. To recover the butyl rubber, the discharge stream is
typically treated with steam and hot water. Most of the solvent and
the unreacted monomers are thereby flashed off. One disadvantage of
the contact with steam and water is, that butyl rubbers are
coagulated. The rubber polymers are then present in the form of wet
crumbs in water. Most of the water is then be separated by
draining, followed e.g. by the application of drying extruders and
a final vacuum drying step.
[0005] In the slurry process, the polymerization reactor discharge
stream is treated with steam and hot water in a flash drum. Most of
the chloromethane and the unreacted monomers are thereby flashed
off and the water is separated from the vapors by condensation.
When the polymer from the reactor is to be processed further, such
as by halogenations, the butyl rubber product may be recovered
directly as a solution by discharging the reactor content into a
hot solvent such as hexane. The chloromethane is evaporated after
this stage and a further stripping stage is applied to remove
remaining monomer residues.
[0006] In the solution process, an inert hydrocarbon solvent and an
aluminium alkyl halide catalyst are applied during the
polymerization step. The remaining monomers are then removed from
the reactor solution in a distillation stripping process. After
this distillation step, the butyl rubber polymer is present as a
homogeneous solution in a hydrocarbon. This solution can either be
processed further, such as being subjected to a halogenation step,
or the butyl rubber polymer can be isolated directly from the
solution. The isolation of the butyl rubber from solution is
similar to that of the slurry process and also involves contact
with steam and hot water, whereby the polymer coagulated. The butyl
rubber polymer is then present in the form of wet crumbs in water
(6 to 10 wt % polymer in water). To counteract the coagulation,
salts of fatty acids are added in the flash drum containing the
butyl rubber crumbs in water following the coagulation/steam
stripping process. After the addition of additives, butyl rubber is
then converted into the final commercial bale form through further
drying. The drying is typically effected by draining, followed by
the application of drying extruders and a final drying step in a
fluidized bed.
[0007] One of the most commercially important chemical modification
of butyl rubber is halogenation which leads to chlorinated and
brominated butyl rubber, hereinafter also denoted as halobutyl
rubbers or individually as bromobutyl rubber or chlorobutyl
rubber.
[0008] Halobutyl rubber is technically produced by contacting a
solution of regular butyl rubber in an alkane with chlorine or
bromine in an agitated vessel. Said solution is generally denoted
as cement. Unreacted halogen and hydrogen halide formed as
byproduct are neutralized by the addition of a caustic solution.
Additives can also be incorporated at that stage. The resulting
solution is then steam-stripped to remove the solvent, thereby
coagulating the rubber into a solid product. The solid product is
generally recovered as a 5 to 12% slurry in water. Stabilizers
and/or antioxidants are added to the halogenated butyl rubber
immediately before recovery. The halogenated butyl rubber is then
finished using mechanical drying equipment in a process analogous
to that used for regular butyl rubber; however, because of the
greater reactivity of the halogenated product, less severe
conditions are employed.
[0009] The aforementioned processes for coagulation and steam
stripping suffer from very high energy consumption. A large amount
of steam is necessary not only to evaporate the solvent but also to
heat and maintain the complete water content of the stripping drums
at a high temperature. Additional steam addition is also necessary
to strip off residual amounts of solvent by lowering the partial
pressure of the solvent in the stripping drum.
[0010] The aforementioned processes also utilize a large amount of
water because the concentration of butyl rubber in the slurry after
coagulation is generally only 5 to 12% by weight and only 5% to 20%
for halogenated butyl rubbers. All water from this slurry
constitutes waste water and must be disposed of. While the waste
water contains sodium salts from the neutralization, reworking and
recycling the waste water to remove the sodium salts is not
economically viable because the salt concentration is too low.
[0011] The rubber crumbs are separated from the bulk water
mechanically using simple sieve trays or screens. The halogenated
butyl rubber still contain approximately 30 to 50% water after this
first separation. Further mechanical drying is then conducted using
extruders by kneading the product and squeezing out the water. The
disadvantage of this mechanical drying process is the contamination
of water by small rubber particles that were not held back by the
sieves with the result that the waste water requires additional
treatment.
[0012] The aforementioned mechanical dewatering can only diminish
moisture content down to approximately 5 to 15%. Additional thermal
drying stages are then required. The halogenated butyl rubber is
thereby heated to 150 to 200.degree. C. under pressure in a single
screw or twin screw extruder. A die plate is installed to maintain
the pressure. When the halogenated butyl rubber is pushed through
the die plate, the water in the rubber evaporates and forms open
porous crumbs. A cutting device then cuts the crumbs into small
pieces. The crumbs are conveyed to a convective dryer where
residual moisture is removed by hot air. After such drying, the
(halo)butyl rubber generally has a moisture content of 0.1 to 0.7%.
A cooling stage, accomplished by flowing cold air through the
rubber crumbs, is then needed to cool the halogenated butyl rubber
crumbs down to the maximum baling temperature of 60.degree. C. The
crumbs are then formed into bales by hydraulic presses, and the
bales are packed into boxes or crates for shipment.
[0013] The aforementioned processes for drying halogenated butyl
rubbers is complex and requires extensive equipment. Furthermore,
the process parameters must be carefully monitored to avoid heat
and shear stress, which would accelerate degradation of the
halogenated butyl rubber.
[0014] Various other special processes have been developed with the
aim of removing water and volatile organic solvents from polymers.
Extruder degassing in vacuum with or without the use of entrainers
has gained acceptance in practical applications as the most
important technique, however, the energy requirements of such prior
art processes are quite high.
[0015] U.S. Pat. No. 3,117,953 A1 discloses an apparatus and
process for purifying high pressure polyethylene. The substitution
of synthetic rubber cement for polyethylene in U.S. Pat. No.
3,117,953 A1 would, however, result in crumbs being formed prior to
entering the extruder, which is not desirable at all.
[0016] DE 195 37 113 discloses a method and an apparatus for
polymer resins in particular polycarbonate resins using a steam
stripper a decanter and an extruder. However, the introduction of
steam would result in an undesirable high content of residual water
or a very high energy consumption.
[0017] U.S. Pat. No. 4,055,001 discloses a method for the
preparation of polymers such as butyl rubber having a water content
of less than 0.1 wt.-% by using ultrasound sonotrodes during the
drying process. However, the very high shear stress associated with
the use of ultrasound is prohibitive for polymers such as halobutyl
rubbers.
[0018] EP 0 102 122 discloses a method for polymer recovery from a
solution, in particular for recovery of polyethylene, using a
partially filled extruder. However, EP 0 102 122 is silent about
the removal of residual water.
[0019] US 2001/056176 A1 discloses a one step method of recovering
a polymer and specifically an example for the concentration of
rubber solutions. The rubber solution is thereby heated with steam
in order to remove existing solvents in one step by degassing under
vacuum to produce white crumb. US 2001/056176 A1 thereby requires a
large volumetric vapor flow to remove the volatile components at
low vapor pressure and results in the enclosure of additional water
in the crumbs, which water would subsequently need to be
removed.
[0020] U.S. Pat. No. 5,283,021 A1 discloses a two step process for
removing solvent from an elastomeric polymer solution. The polymer
solution is thereby heated directly by a heating fluid and sprayed
under vacuum. During the spraying, the solvent is evaporated,
thereby forming crumbs which are then fed to an extruder for
further degassing. However, crumb formation at that stage is not
desirable.
[0021] EP 1 127 609 A2 discloses a process to treat a product in at
least one kneader. EP 1 127 609 A2 uses energy introduced in part
through the wall of the kneader itself to evaporate the solvent
from solutions containing elastomers and thermoplastics. A kneader
with a large surface area is therefore required as are high
investment costs. Another portion of the energy is introduced via
the rotating shaft of the kneader as mechanical energy. Mechanical
energy is more expensive and therefore environmentally
disadvantageous when compared to steam heating. The kneaders used
in EP 1 127 609 A2 require a great deal of maintenance and
cleaning. The introduction of mechanical energy via the kneader is
furthermore strongly dependent on the viscosity of the product,
which reduces the flexibility of the process.
[0022] EP 1 165 302 A 1 discloses a device and method for degassing
plastics. The apparatus in EP 1 165 302 A1 is an extruder with a
rear vent and several vent sections operated under vacuum. The
vacuum is needed to achieve low residual volatile concentrations.
EP 1 165 302 A1 discloses that a stripping agent can be applied to
further improve degassing efficiency. The plastic used in EP 1 165
302 A1, the thermoplastic polycarbonate, remains a flowing melt at
the end of the degassing process. A synthetic rubber cement
processed pursuant to EP 1 165 302 A1 would, however, convert to
crumbs at the end of the degassing stage and could not be processed
further.
[0023] In "Process Machinery", Parts I and II, March and April
2000; Author: C. G. Hagberg, a direct volatilization of rubber
solutions using a flash tank and an extruder is disclosed. However,
this reference is silent about the contents of volatile compounds
in the final product.
[0024] PCT/EP2009/062073 discloses a device and method for
degassing non-volatile polymers. The device preferably comprises a
twin screw extruder with a rear vent and several forward directed
vent sections.
[0025] In view of the foregoing, an object of the present invention
was therefore to provide a continuous, energy efficient,
ecologically and economically favourable process to remove volatile
compounds from a fluid containing at least one halogenated butyl
rubber producing a halogenated butyl rubber product that is
substantially free of volatile compounds.
[0026] This object is solved by a process of removing volatile
compounds from a fluid containing at least one halogenated butyl
rubber and at least one volatile compound which comprises at least
the steps of: [0027] a) treating the fluid in at least one
concentrator unit comprising at least a heater, a degassing vessel
and a vapor line, whereby the fluid is heated, the heated fluid is
fed into a degassing vessel where part of the volatile compounds
are removed via the vapor line to obtain a concentrated fluid,
[0028] b) reheating the concentrated fluid from step a) in at least
one reheating unit to obtain a reheated concentrated fluid; [0029]
c) feeding the reheated concentrated fluid from step b) into at
least one extruder unit comprising at least an extruder degassing
section comprising at least a conveying section, a vent port with
one or more vapor lines, a accumulating section and an outlet
section, whereby volatile compounds are removed through the vent
ports and vapor lines; whereby the product obtained at the outlet
section is substantially free of volatile compounds.
[0030] It is pointed out that the scope of the invention also
encompasses any desired combinations of the ranges and areas of
preference specified for each feature.
[0031] In a preferred embodiment of the invention, the reheated
concentrated fluid (L) is free-flowing. In the context of this
invention, the term "free-flowing" means a viscosity in the range
of 100 to 50,000,000 mPa*s, preferably 5,000 to 30,000,000 mPa*s
and most preferably 10,000 mPa*s to 300,000 mPa*s.
[0032] As far as not mentioned otherwise the viscosity values of
fluids refer to the zero shear viscosity extrapolated from
measurements at given temperature using a Haake Rheostress RS 150
viscosimeter or a rotational rheometer of cone-plate type for very
viscuous samples. The extrapolation is performed by taking a
2.sup.nd order polynomial to reflect the shear stress vs shear rate
graph obtained from the measurements. The linear portion of the
polynomial reflects the slope at a shear rate of zero and thus is
the zero shear viscosity.
[0033] In the context of this invention, the term "substantially
free of volatile compounds" means a total concentration of volatile
compounds of less than 1 wt %, preferably less than 0.5 wt % based
on the mass of the non-volatile polymer.
[0034] In particular, the term "substantially free of volatile
compounds" means substantially free of water and substantially free
of volatile organic compounds.
[0035] Non-volatile polymers are considered to be substantially
free of water, if the residual water concentration is less than 0.5
wt % preferably less than 0.25 wt %, more preferably less than 0.1
wt % and most preferably less than 0.075 wt % based on the mass of
the polymer.
[0036] In the context of this invention, the term "volatile organic
compounds" means organic compounds having a boiling point of below
250.degree. C. at standard pressure.
[0037] Halogenated butyl rubbers are considered substantially free
of volatile organic compound, if the residual concentration of said
volatile organic compounds is less than 0.75 wt % preferably less
than 0.25 wt % and most preferably less than 0.1 wt % based on the
mass of the halogenated butyl rubber. Said volatile organic
compounds are typically the solvents employed in the halogenation
step following polymerization and include hydrocarbons like hexanes
and pentanes.
[0038] As used herein, the term halogenated butyl rubber includes
bromo- and chlorobutyl rubbers, brominated and/or chlorinated
terpolymers such as those described in U.S. Pat. No. 6,960,632 and
Kaszas et al., Rubber Chemistry and Technology, 2001, 75, 155 where
para-methylstyrene is added to the mixed feed of butyl
polymerizations (Methyl chloride, isobutylene and isoprene mixed
feed, with aluminum trichloride/water mixtures as initiator)
resulting in a high molecular weight polymer with up to 10 mol % of
styrenic groups randomly incorporated along the polymer chain The
incorporation of para-methylstyrene is found to be uniform
throughout the molecular weight distribution due to the similarity
in reactivity with isobutylene. The isoprene moieties within the
butyl terpolymers can be brominated by conventional methods.
Alternatively, a brominated and/or chlorinated terpolymer may
comprise a C.sub.4 to C.sub.7 isomonoolefin, such as isobutylene,
and a comonomer, such as para-alkylstyrene, preferably
para-methylstrene. The aforementioned copolymers are commercially
available under the tradename EXXPRO 3035, 3433, 3745. When
halogenated, some of the alkyl substituent groups present in the
styrene monomer units contain a benzylic halide formed from
halogenation of the polymer.
[0039] In the context of this invention butyl rubber preferably
denotes a (co)-polymer of isobutene (2-methylpropene) and isoprene
(2-methylbuta-1,3-diene). On a molar basis, the isoprene content in
the polymer is between 0.001% and 20, preferably between 1.8 and
2.3 mol %. Butyl rubber is composed of linear polyisobutene chains
with randomly distributed isoprene units. The isoprene units
introduce unsaturated sites into the polymer chain to enable
vulcanization. The mass average molecular weight of butyl rubber
molecules Mw is typically between 50,000 and 1,000,000 g/mol,
preferably between 300,000 and 1,000,000 g/mol.
[0040] The halogenated butyl rubbers also contain a certain amount
of halogen chemically bound to the rubber molecules. The amount of
chemically bound halogen is typically in the range of more than 0
to 8 wt % with respect to total mass of the polymer. The
halogenated butyl rubbers may also contain additives, e.g. 0.0001
to 4 phr (phr=parts per hundred rubber with respect to rubber
weight), epoxidized soy bean oil (ESBO), 0.0001 to 5 phr
calcium-stearate and 0.0001 to 0.5 phr antioxidants. Other
additives are also applicable, dependent on the application of the
butyl rubber product, i.e. fillers or colorants.
[0041] In case of bromobutyl rubber, the typical bromine content in
the finished product is 1.5 to 5 wt %, preferably 1.6 to 2.0 wt
%.
[0042] In case of chlorobutyl rubber, the typical chlorine content
in the finished product is 1.0 to 5 wt %, preferably 1.15 to 1.35
wt %.
[0043] The subject of the invention will be described in more
detail by means of schematic drawings in which:
[0044] FIG. 1 shows a single-stage concentrator unit, a reheating
unit and an extruder unit comprising one extruder degassing
section, one accumulating section and one outlet section.
[0045] FIG. 2 shows a single-stage concentrator unit, a reheating
unit and an extruder unit comprising two extruder degassing
sections, two accumulating sections and one outlet section.
[0046] FIG. 3 shows a single-stage concentrator unit having a
pressure relief valve, a reheating unit and an extruder unit having
a pressure relief valve and further comprising two extruder
degassing sections, two accumulating sections, a side feeder and an
outlet section.
[0047] FIG. 4 shows a double-stage concentrator unit, a reheating
unit and an extruder unit comprising one extruder degassing
section, one accumulating section and an outlet section.
[0048] FIG. 5 shows a single-stage concentrator unit, a reheating
unit and an extruder unit comprising three extruder degassing
sections, three accumulating sections and one outlet section,
whereby one extruder degassing section is a backward degassing
section.
[0049] FIG. 6 shows a single-stage concentrator unit comprising a
pressure regulation device, a reheating unit and an extruder unit
comprising a pressure regulation device, four extruder degassing
sections, four accumulating sections and one outlet section,
whereby one extruder degassing section is a backward degassing
section.
[0050] FIG. 7 shows a single-stage prewashing unit, a single-stage
concentrator unit, a reheating unit and an extruder unit comprising
one extruder degassing section, one accumulating section and one
outlet section.
[0051] FIG. 8 shows a basic prewashing unit
[0052] FIG. 9 shows a prewashing unit comprising a coalescer
[0053] FIG. 10 shows a double-stage prewashing unit
[0054] FIG. 11 shows a double-stage prewashing unit having
additional heaters
[0055] FIG. 12 shows a flow chart for a preferred process.
[0056] A basic and exemplary embodiment of the process step is
shown in FIG. 1. In step a) Fluid F containing at least one
halogenated butyl rubber and at least one volatile compound is
transferred via pump 1 to the heater 2, where the fluid F is
heated.
[0057] Fluid F, also called cement, contains for example from 3 to
50 wt % of a halogenated butyl rubber and from 60 to 97 wt %
volatile compounds, in particular a solvent or a solvent and water,
whereby the aforementioned components add up to 90 to 100,
preferably 95 to 100 wt % of the total mass of fluid F.
[0058] The solvent is preferably selected from the group consisting
of linear or branched alkanes having between 4 and 10 C atoms,
preferably 4 to 7 C atoms. More preferred solvents are n-pentane,
iso-pentane, n-hexane, cyclo-hexane, iso-hexane,
methyl-cyclopentane, methyl-cyclohexane and n-heptane as well as
mixtures comprising or consisting of those alkanes.
[0059] In a preferred embodiment of the invention, fluid F contains
from 3 to 40 wt % of a halogenated butyl rubber from 60 to 95 wt %
volatile organic compounds, in particular a solvent, and from 0.5
to 20 wt % water, whereby the aforementioned components add up to
95 to 100 wt % of the total mass of fluid F.
[0060] The fluid F is typically obtained from halogenation of butyl
rubber. Fluids F containing water are typically obtained after
steam stripping processes following the polymerization.
[0061] The fluid F entering the heater typically and preferably has
a temperature of 10.degree. C. to 100.degree. C., preferably of
30.degree. C. to 80.degree. C. The viscosity of fluid F is for
example in the range of 100 mPa*s to 25,000 mPa*s, preferably in
the range of 500 mPa*s to 5,000 mPa*s.
[0062] A heater may be any device that is able to raise the
temperature of Fluid F. In a preferred embodiment, heater 2 is a
heat exchanger. The heating medium is selected from the group
consisting of steam, heating oil or hot pressurized water. The heat
exchanger is for example of shell-and-tube type, where the fluid F
is inside the tubes and the heating medium is on the shell side.
Special inserts in the tubes may be applied to enhance heat
transfer. Another type of heat exchanger may also be used, in which
fluid F is on the outside of the heat exchanger tubes. The
advantage of the aforementioned types of heat exchangers is the
avoidance of maldistribution and easy maintenance as well as good
heat transfer. Said heat exchangers are well known and commercially
available. In a less preferred embodiment Plate type heat
exchangers may also be applied.
[0063] Upon heating, heated fluid G is obtained. The heated fluid G
has a higher temperature than fluid F, preferably a temperature of
100 to 200.degree. C., more preferably 110.degree. C. to
190.degree. C. and even more preferably 120.degree. C. to
175.degree. C. The heated fluid G is then conveyed further into a
degassing vessel 4. In the degassing vessel, the volatile compounds
at least partially evaporate. The vapors are separated and removed
from the heated fluid G by a vacuum line 4.1. The pressure in the
degassing vessel 4 is for example in the range of 100 hPa to 4,000
hPa, preferably in the range of 200 hPa and 2,000 hPa and more
preferred in the range of 230 to 1,100 hPa.
[0064] The vapors removed via the vacuum line 4.1 are preferably
condensed and recycled into the process for preparation of fluid F.
After degassing and separation a concentrated fluid H is obtained,
which is removed from the degassing vessel 4 by means of a pump
4.2.
[0065] Generally the degassing vessel may be a flash evaporator or
another device typically used to remove volatile compounds while
simultaneously having short retention times.
[0066] In a preferred embodiment of the invention the degassing
vessel is designed in the shape of a cyclone to further aid
separation of vapor from heated fluid G. In another preferred
embodiment of the invention the degassing vessel 4 has a conical or
at least torisperical shaped bottom, to allow the vessel being
emptied completely or substantially complete.
[0067] In another embodiment the inner surface of the degassing
vessel can be heated.
[0068] The pump 4.2 is preferably directly connected to the outlet
of the degassing vessel 4. In general, the connection piece between
pump and vessel is preferably as short as possible.
[0069] Due to the high viscosity of the concentrated fluid H at
this stage, the inlet of the pump is preferably designed with a
large inlet, thereby reducing the pressure drop at the inlet. The
inlet comprises a cross sectional area A.sub.inlet and the
degassing vessel comprises an inner surface A.sub.degas, wherein
the ratio A.sub.inlet/A.sub.degas is particularly
0.001.ltoreq.A.sub.inlet/A.sub.degas.ltoreq.0.4, preferably
0.01.ltoreq.A.sub.inlet/A.sub.degas.ltoreq.0.3, more preferred
0.05.ltoreq.A.sub.inlet/A.sub.degas.ltoreq.0.25 and most preferred
0.1.ltoreq.A.sub.inlet/A.sub.degas.ltoreq.0.2.
[0070] The pump 4.2 may be selected from the group consisting of
positive displacement type pumps, gear pumps, piston pumps,
membrane pumps, screw type pumps, extruder type pumps like
counter-rotating or co-rotating single or twin screw extruders or
kneader type pumps. Positive displacement type pumps and gear pumps
are preferred, gear pumps are even more preferred.
[0071] In another preferred embodiment the pump 4.2 comprises a
combination of an extruder or a kneader and a gear pump whereby the
gear pump is fed from the extruder or kneader.
[0072] The amount of volatile compounds that is removed in this
step a) is for example dependent on the temperature of fluid G and
the pressure in the degassing vessel 4. In a preferred embodiment
of the invention the temperature of fluid G and the pressure in the
degassing vessel 4 are chosen so that the concentrated fluid H is
preferably free-flowing as defined above and comprises for example
from 10 to 60, preferably from 25 to 60 wt % of a non-volatile
polymer, preferably a synthetic rubber and more preferably
(halo)butyl rubber and from about 40 to about 90, preferably from
40 to 75 wt % volatile compounds whereby the aforementioned
components non-volatile polymer, volatile organic compound and
water add up to 90 to 100 wt %, preferably to 95 to 100 wt % of the
total mass of fluid H.
[0073] In a preferred embodiment and where the feedstock fluid F
comprises water, fluid H for example comprises from 10 to 60,
preferably from 25 to 60 wt % of a non-volatile polymer, preferably
a synthetic rubber and more preferably (halo)butyl rubber, from
about 25 to about 90, preferably from 25 to 75 wt % volatile
organic compounds, in particular a solvent, and about 0.5 to about
15 wt % water, whereby the aforementioned components non-volatile
polymer, volatile organic compound and water add up to 90 to 100 wt
%, preferably 95 to 100 wt % of the total mass of fluid H.
[0074] The temperature of the concentrated fluid H is lower than
that of heated fluid G and is for example in the range of 15 to
100.degree. C., preferably in the range of 30 to 100.degree. C. The
pressure of the heated fluid G is for example in the range of 2 to
60 bar, preferably in the range of 4 to 30 bar. The concentrated
fluid H is preferably free-flowing as defined above.
[0075] In step b), the concentrated fluid H obtained in step a) is
then passed through a reheating unit 6 to obtain a reheated
concentrated fluid L. The a preferred embodiment the reheating unit
comprises a heat exchanger, whereby the same disclosure including
the preferences with regard to heating media and heat exchanger
types apply as described above for heat exchanger 2.
[0076] The temperature of the reheated concentrated fluid L is
higher than that of the concentrated fluid L and is for example in
the range 50.degree. C. to 200.degree. C., preferably in the range
of 90.degree. C. to 180.degree. C. The pressure of the heated fluid
G is for example in the range of 2 to 60 bar, preferably in the
range of 4 to 30 bar. The reheated concentrated fluid L is
preferably free-flowing as defined above.
[0077] In step c), the reheated concentrated fluid L obtained in
step b) is passed on to a extruder unit and fed into the conveying
section 16 of the extruder degassing section at the feeding point
12.
[0078] Suitable extruder types include single screw and multiscrew
extruders comprising any number of barrels and types of screw
elements and other single or multishaft conveying kneaders.
Possible embodiments of multiscrew extruders are twin-screw
extruders, ring extruders or planetary roller extruders, whereby
twin-screw extruders and planetary roller extruders are
preferred.
[0079] Single screw extruders include those having an axial
oscillating screw. Twin screw extruders are for example
counter-rotating intermeshing, counter-rotating non-intermeshing,
co-rotating intermeshing and co-rotating non-intermeshing twin
screw extruders, whereby co-rotating intermeshing twin screw
extruders are preferred.
[0080] In one embodiment of the invention the extruders can either
be heated via the barrels to temperatures up to 300.degree. C. or
cooled.
[0081] In a preferred embodiment, the extruder comprises means to
operate separate zones independently of each other at different
temperatures so that the zones can either be heated, unheated or
cooled. In another preferred embodiment the extruder comprises for
each conveying section at least one separate zone, which can be
operated independently at different temperatures.
[0082] Preferred extruder materials should be non-corrosive and
should substantially prevent the reheated concentrated fluid L and
the Product P from being contaminated with metal or metal ions.
[0083] Preferred extruder materials include nitrided steel, duplex
steel, stainless steel, nickel-based alloys, composite materials
like sintered metals, hot isostatic pressed materials, hard wear
resistant materials like Stellite, coated metals with coatings for
example made from ceramics, titanium nitride, chromium nitride and
diamond like carbon (DLC).
[0084] The conveying section 16 is open to a vent port 15. In the
conveying section 16 a part of the solvent is evaporated and
separated from the reheated concentrated fluid L. The vapors are
removed through the vent port 15 via a vapor line 15.1.
[0085] Since the evaporation volatile compounds have a tendency to
entrain the reheated concentrated fluid L or the product P towards
the vent ports, in a preferred embodiment of the invention the vent
ports 15 are designed to prevent the material, in particular the
reheated concentrated fluid L or the Product P, from coming out of
the vent ports.
[0086] Suitable means to accomplish that purpose are stuffer
screws, that are mounted on the vent ports and convey any material
back into the extruder, or rollers or belts, that are applied to
the inside of the vent ports to push deposited material back into
the extruder.
[0087] Stuffer screws are preferred. The stuffer screws may
comprise one two or more shafts, whereby stuffer screws comprising
one or two shafts are preferred.
[0088] As an alternative or in addition to the aforementioned,
coatings of the vent ports may be applied which reduce or prevent
sticking of the material to the surface. Suitable coatings include
DLC, Ethylene-Tetrafluoroethylene (ETFE), Polytetrafluoroethylene
(PTFE) and Nickel-Alloys.
[0089] The pressure at the vent port 15 is for example between 1
hPa and 2,000 hPa, preferably between 5 hPa and 900 hPa.
[0090] The vapor line 15.1 may be and is preferably connected to a
condensing system.
[0091] In general, the purpose of the condensing system is to
collect volatile compounds removed by the vent ports via the vapour
lines and typically comprises a condenser and a vacuum pump. Any
condensing system known in the art may be used to effect the
recovery of volatile compounds.
[0092] Generally, it is preferred to recycle the condensed volatile
compounds, optionally after carrying out a phase separation to
separate the volatile organic compounds from water, into a process
for the preparation of fluid F or A as defined below.
[0093] The conveying section 16 is terminated by a accumulating
section 20. The purpose of the accumulation is to assure a certain
pressure level in the vent port 15 and to introduce mechanical
energy into the material to facilitate evaporation of volatile
compounds. The accumulating section 20 may comprise any means that
enable the accumulation of the material. It may be designed to
include for example kneading or throttling elements, blister discs
or die plates.
[0094] Examples of throttling elements are conical or cylindrical
flow paths or other throttling means.
[0095] The application of kneading elements, blister discs or die
plates within the accumulating section is preferred, kneading
elements are even more preferred. Examples of kneading elements
include kneading blocks, which may be designed as double or triple
flighted forward, backward or neutral conveying kneading blocks;
single or double flighted screw mixing elements with grooves,
single flighted tooth mixing elements, blister plates and single,
double or triple flighted eccentric discs. The kneading elements
may be assembled in any combination on the screw shafts of the
extruder, in particular of an twin screw counter rotating or
co-rotating twin screw extruder.
[0096] A typical accumulating section comprises of 2 to 10 kneading
blocks, oftentimes terminated by a back conveying type of kneading
element. For mixing in of a stripping agent, tooth type elements or
screw elements with grooves may be applied.
[0097] Eccentric discs are preferably applied in the last section
of the extruder, where the product P is highly viscous and
substantially free of volatile compounds
[0098] For planetary roller extruders, kneading elements like tooth
shaped rollers are or rollers with grooves and clearances are
preferred.
[0099] Generally the extruder unit may comprise one or more
conveying sections and one or more accumulating sections, whereby
the number is only limited by constructional constraints. A typical
number of conveying sections and accumulating sections is 1 to 30,
preferably 2 to 20 and more preferably 3 to 15.
[0100] The last accumulating section 20 is typically designed to
form a product plug at the outlet of the extruder, thereby
preventing surrounding air from entering the extruder. While
passing from the conveying section 16 and the accumulating section
20 to the outlet section 22 the reheated concentrated fluid L
undergoes a transition from the preferably free-flowing reheated
concentrated fluid L to the product P, which typically has a
crumbly appearance.
[0101] The outlet section 22 typically comprises means to allow the
product to exit the extruder and optionally but preferably product
processing equipment. Examples of suitable product processing
equipment includes combinations of die plates and cutters; die
plates and underwater-pelletizing means; means for crumb formation
like screw elements with teeth and holes; turbulators which may be
designed as cylinders with holes in it, whereby the product is
pressed from the outside to the inside of the cylinder, and whereby
a rotating knife inside the cylinder cuts the product into pieces;
fixed knifes placed at the end plate of the extruder, whereby the
screw rotation causes the cutting action, which preferably is
applied when working with twin screw co-rotating, single screw and
planetary roller extruders.
[0102] To reduce the mechanical and thermal stress to the product,
in a preferred embodiment of the invention the product processing
equipment is combined with cooling means.
[0103] The cooling means comprises any means that allow the removal
of heat from the product. Examples of cooling means include
pneumatic crumb conveyers with convective air cooling, vibrating
crumb conveyers with convective air cooling, vibrating crumb
conveyer with cooled contact surfaces, belt conveyer with
convective air cooling, belt conveyer with cooled belts, water
spraying on hot crumbs upon outlet of the extruder and as already
mentioned underwater-pelletizing means, whereby water serves as the
coolant.
[0104] The product P may then be processed further for final
packing and shipping. Halogenated butyl rubbers are typically
cooled to a temperature of or below 60.degree. C., formed into
bales e.g. by a hydraulic press, and then packed into boxes or
crates for shipment.
[0105] In general, an increasing feed rate of the reheated
concentrated fluid L at the feeding point 12 requires a
corresponding increase in the screw speed of the extruder.
Moreover, the screw speed determines the residence time of fluid L.
Thus, the screw speed, feed rate and the extruder diameter are
typically interdependent. Typically the extruder is operated in
such a manner that the dimensionless throughput V/(n*d.sup.3),
wherein V denotes the Volume flow rate, n the screw speed expressed
in revolutions per minute and d the effective diameter of the
extruder is adjusted to about 0.01 to about 0.2 preferably to about
0.015 to about 0.1. The maximum and minimum feed rates and extruder
screw speeds are determined by for example the size of the
extruder, the physical properties of the halogenated butyl rubber
contained in Fluid L and the target values of remaining volatile
compounds. Given these properties, however, the operating
parameters can be determined by one skilled in the art by some
initial experiments.
[0106] In one embodiment of the invention the extruder is operated
at a feed rate of 5 to 25,000, preferably of 5 to 6,000 kilograms
per hour.
[0107] Generally, the degassing in the extruder may be aided by the
addition of a stripping agent that is removed together with other
volatile compounds. Even though the stripping agent may be added
anywhere in the extruder unit, the addition in one or more
accumulating sections is preferred. In a more preferred embodiment
a stripping agent is added in one or more accumulating sections
except the last one 20.
[0108] Suitable stripping agents are substances that are inert to
the reheated concentrated fluid L and/or the product P and have a
vapor pressure greater than 100 hPa at 100.degree. C.
[0109] In the context of the invention, the term "inert" means that
the stripping agent does not or virtually not react with the
polymers contained in the reheated concentrated fluid L and/or the
product P. Suitable stripping agents are nitrogen, carbon dioxide,
noble gases, propane, butane, water or a mixture of the
aforementioned substances, whereby carbon dioxide is preferred. The
amount of stripping agent may be 0.0001 to 10, preferably 0.001 to
5 and more preferably 0.1 to 2 wt-% based on the amount of the
polymer product obtained at the outlet section.
[0110] The heating stream of the heating unit 6 may be used after
heating the concentrated fluid H for heating the fluid F in the
heater 2. The heating stream of the reheating unit 6 may be in
communication with the heater 2. In addition or in alternate the
heating stream leaving the heating unit 6 and/or the heating stream
entering the reheating unit 6 may be in communication with a
further reheating unit 6 and/or a further heater 2 as illustrated
in FIG. 4. Preferably the heating stream leaving the reheating unit
6 and/or the heating stream entering the reheating unit 6 may be in
communication with one or more degassing vessels 4 and/or in
communication with one or more extruder units. Further it is
possible that the heating stream leaving the heater 2 and/or the
heating stream entering the heater 2 may be in communication with
one or more degassing vessel 4 and/or in communication with one or
more extruder units. Particularly preferred the heating stream of
the heater 2 and/or of the reheater unit 6 are led in counter flow
with respect to the heated fluids. Due to a suitable connection of
the heating streams of the heater 2, the reheating unit 6 and if so
the degassing vessel 2 and/or the extruder unit a large amount of
the heat content of the heating stream can be used. This leads to
an increased energy efficiency with respect to the required heat
flows at different devices. If necessary, the heating stream may be
heated additionally between two different devices for controlling a
required temperature of the heating stream. In most cases this
additional heating of the heating stream may take place at lower
temperatures and at a lower exergy level compared to the
environment so that the additional heating of the heating stream
can be facilitated and enables a better efficiency.
[0111] Another embodiment of the invention is shown in FIG. 2. FIG.
2 shows another flow chart and suitable device for the
accomplishment of the process according to the invention comprising
a concentrator unit with a pump 1, a heater 2, a degassing vessel
4, a vapour line 4.1 and a pump 4.2, a reheating unit comprising a
reheater unit 6 and an extruder unit comprising two extruder
degassing sections having two conveying sections 16A and 16B each
connected to a vent port 15 A and 15 B and a vapour line 15.1A and
15.1.B, two accumulating sections 18 and 20 terminating the
conveying sections 16 A and 16 B a an outlet section 22. In
addition to that the extruder unit further comprises a side feeder
24.
[0112] Generally, the extruder unit may comprise one or more side
feeders, which may positioned anywhere in the extruder, preferably
in close proximity to the feeding point or the outlet section 22.
Side feeders are suitable for the addition of additives to the
polymer.
[0113] Examples of additives halogenated butyl rubber products
include stabilizing agents, acid scavengers like ESBO (epoxidized
soy bean oil), stearates like calcium stearates, antioxidants and
the like. Examples of suitable antioxidants include sterically
hindered phenols like butylhydroxytoluenes and its derivatives like
Inganox 1010 and 1076, amines, mercapto-benzimidazoles, certain
phosphites and the like.
[0114] In particular, halogenated butyl rubber are mixed with
additives, e.g. 0.0001 to 4 phr epoxidized soy bean oil (ESBO),
0.0001 to 5 phr calcium-stearate and 0.0001 to 0.5 phr of
antioxidants (phr=parts per hundred rubber with respect to rubber
weight). Other additives are also applicable, dependent on the
application of the butyl rubber product, i.e. fillers or
colorants.
[0115] As an alternative or in addition to that, additives may also
already be added to the fluid F or, as far as they are liquid
together with the stripping agent.
[0116] In a preferred embodiment of the invention step a) is
repeated a least once, preferably once or twice. The advantage of
repeating step a) is that the total energy consumption to produce
the concentrated fluid H can significantly reduced due to easier
operation parameter optimization for each concentration unit. The
repetition of step a) is preferably accomplished by connecting the
respective number of concentrating units in series.
[0117] An example of this embodiment is shown in FIG. 4. FIG. 4
shows another flow chart and suitable device for the accomplishment
of the process according to the invention comprising a double-stage
concentrator unit with a pump 1, a first concentrator unit
comprising heater 2A, degassing vessel 4A equipped with a vapour
line 4.1A and a pump 4.2A, a second concentrator unit comprising
heater 2B, degassing vessel 4B equipped with a vapour line 4.1B and
a pump 4.2B, a reheating unit 6 and an extruder unit connected to a
vent port 15 A and a vapour line 15.1A. The heated fluid G is
subjected to the first concentration stage, thereby obtaining
pre-concentrated fluid J, which is then reheated by heater 2B to
obtain the reheated pre-concentrated fluid K, which is then
subjected to the second concentration stage, whereby concentrated
fluid H is obtained. Concentrated fluid H is then processed further
as described above.
[0118] In a preferred embodiment of the invention the concentration
unit, the reheating unit or the extruder unit may independently of
each other be equipped with one or more pressure regulation devices
which allow the very precise operation of the units under
predefined conditions.
[0119] The pressure regulation devices may be active or passive,
whereby active pressure regulation devices are preferred. Examples
of active pressure regulation devices include control valves like a
pressure relief valve, examples of passive pressure regulation
devices include nozzles and dies or orifice plates. Suitable valves
may be selected from ball, piston, gate or needle valves.
[0120] In case of a passive pressure control device, it is
preferred to calculate an orifice to cause a certain pressure drop.
The calculation is based on viscosity of the fluid at that point
and the throughput. Anyone skilled in the art can perform this
calculation.
[0121] Active pressure control devices are typically controlled by
a pressure measurement upstream of the device. The pressure is for
example measured and compared to the set point. The pressure
control device is then adjusted according to the offset
recognized.
[0122] Alternatively the pressure drop across the device is
measured instead of the absolute pressure upstream of the pressure
control device. The valve position is adjusted manually,
electrically, pneumatically or hydraulically. The control of the
valve position, i.e. adjustment to the set point pressure, can for
example be made manually or from any automated process control
system.
[0123] A further embodiment of the invention having additional
pressure control devices is shown in FIG. 3 which is apart form the
pressure control devices very similar to FIG. 2. The pressure of
heated fluid G is controlled by the pressure control device 3, the
pressure of reheated, concentrated fluid L entering the extruder is
controlled by the pressure control device 7.
[0124] In a preferred embodiment of the invention the reheated
concentrated fluid L is injected into the first extruder degassing
section of the extruder unit, whereby the first extruder degassing
section comprises one or more rear vent ports in upstream direction
each connected to a vapor line.
[0125] The advantage of rear vent ports is that the volatile
compounds present in the reheated concentrated fluid L undergo
sudden and rapid evaporation, thereby effecting at least partial
separation of the synthetic rubber product and the volatile
compounds, the vapors emerging through the rear vents in upstream
direction. Generally, from about 50 to about 99 wt-%, of the
volatile compounds present in the fluid L is removed through the
upstream vents.
[0126] An example of this embodiment is shown in FIG. 5. FIG. 5
shows another flow chart and suitable device for the accomplishment
of the process according to the invention comprising a single-stage
concentrator unit with a pump 1, a concentrator unit comprising
heater 2, degassing vessel 4 equipped with a vapour line 4.1 and a
pump 4.2, a reheating unit 6 and an extruder unit comprising three
extruder degassing sections, whereby the feeding point 12 is
located at the first extruder degassing section, comprising a
conveying section 16A, a rear vent port 13 connected to a vapor
line 13.1 in upstream direction and whereby the extruder unit
further comprises two downstream extruder degassing sections each
comprising a conveying section 16 B and 16 C, a vent port, 15 A and
15B, whereby the vent ports 15A and 15B are each connected to a
vapour line 15.1A and 15.1B, and whereby each of the conveying
sections 16A, 16B and 16C is terminated by a accumulating section
18A, 18B and 20 and whereby the extruder unit further comprises an
outlet section 22. Generally the streams are processed as described
above with the difference being that large amounts of fluid
compounds present in the reheated concentrated fluid L are already
removed via vent port 13 and the vapour line 13.1 connected
thereto.
[0127] Another example of this embodiment is shown in FIG. 6. FIG.
6 shows another flow chart and suitable device for the
accomplishment of the process according to the invention comprising
a single-stage concentrator unit with a pump 1, a concentrator unit
comprising a pressure control device 3, a heater 2, a degassing
vessel 4 equipped with a vapour line 4.1 and a pump 4.2, a
reheating unit comprising a heater 6 and an extruder unit
comprising a pressure control device 7 upstream the feeding point
12 of the extruder, four extruder degassing sections, whereby the
feeding point 12 is located at the first extruder degassing
section, whereby the first extruder degassing section comprises a
conveying section 16A, a rear vent port 13 connected to a vapor
line 13.1 in upstream direction and whereby the extruder unit
further comprises three downstream extruder degassing sections each
comprising a conveying section, 16 B, 16 C and 16D, a vent port,
15A, 15B and 15C, whereby the vent ports 15A, 15B and 15C are each
connected to a vapour line 15.1A, 15.1B and 15C, and whereby each
of the conveying sections 16A, 16B, 16C and 16D is terminated by a
accumulating section 18A, 18B, 18C and 20 and whereby the extruder
unit further comprises an outlet section 22. Generally, the streams
are processed as described above.
[0128] Fluid F, which is fed into the heater 2 typically, and as
already disclosed above, contains for example from 3 to 50 wt % of
a non-volatile polymer, preferably a synthetic rubber and more
preferably a (halo)butyl rubber and from 60 to 97 wt % volatile
compounds, in particular a solvent or a solvent and water, whereby
the aforementioned components add up to 90 to 100, preferably 95 to
100 wt % of the total mass of fluid F and in a preferred embodiment
from 3 to 40 wt % of a non-volatile polymer, preferably a synthetic
rubber and more preferably (halo)butyl rubber, from 60 to 95 wt %
volatile organic compounds, in particular a solvent, and from 0.5
to 20 wt % water, whereby the aforementioned components add up to
95 to 100 wt % of the total mass of fluid F.
[0129] Dependant on the source of fluid F it further may contain
hydrophilic compounds which need to be removed to a certain extend
in order to met the desired product specifications.
[0130] Furthermore, where fluid F contains water, it is desirable
to lower the water content in order to improve the process with
respect to its energy consumption.
[0131] It was found that a significant reduction of remaining
hydrophilic compounds or water or both can be achieved in an
advantageous way by preparing the fluid F in a process of removing
hydrophilic compounds and optionally water from a crude fluid A
containing at least one non-volatile polymer, at least one volatile
organic compound, one or more hydrophilic compounds and optionally
water which comprises at least the step of [0132] pre a) treating
the crude fluid A in at least one pre-washing unit comprising at
least a separating apparatus 26, whereby the fluid A is mixed with
water to obtain an organic phase 28 comprising primarily
non-volatile polymer and volatile organic compounds and an aqueous
phase 27 comprising primarily water and hydrophilic compounds, and
whereby the organic phase 28 is separated from the aqueous phase 27
in a separating apparatus 26 and further used as fluid F and
whereby at least a part of the aqueous phase 27 is removed from the
separating apparatus (fluid C).
[0133] In the context of this invention the term "hydrophilic
compounds" denotes at least partially water-soluble volatile and
non-volatile compounds. Examples include inorganic salts and in
particular residues of catalysts employed for the polymerization
reaction like e.g. aluminum salts, iron or other transition metal
salts or halides resulting from halogenation reactions and
neutralizations.
[0134] Exemplary embodiments of step pre-a) are illustrated using
FIGS. 8, 9, 10 and 11.
[0135] A very basic and exemplary embodiment of the pre-washing
step is shown in FIG. 8. In step pre-a) Fluid A containing at least
one non-volatile polymer, at least one volatile compound and at
least one hydrophilic compound is transferred to the separating
apparatus 26, where it is mixed with water. Upon mixing with water
an organic phase 28 and an aqueous phase 27 are obtained. The
organic phase 28 is removed from the separating apparatus 26 and
further used as fluid F, the aqueous phase 27 is at least partially
removed from the separating apparatus 26 as fluid C, which is
disposed of.
[0136] An improved embodiment of the pre-washing step is shown in
FIG. 9. In step pre-a) crude fluid A containing at least one
non-volatile polymer, at least one volatile compound and at least
one hydrophilic compound is fed to the mixing section 30 of the
separating apparatus 26, which is equipped with a mixer 32 and
passes through the separating wall 34 into a settling section,
where the mixture separates into an aqueous phase 27 and an organic
phase 28, whereby the separation is supported by means of a
coalescer 39. A part of the aqueous phase 27 is removed from the
separating apparatus 26 as fluid C, which is typically disposed of,
with the rest being enriched with fresh water E and recycled via
the recirculation line 38 by the action of recirculation pump 36
back into the mixing section 30. The organic phase 28 is removed
and subjected to the subsequent process according to steps a) to c)
as fluid F.
[0137] Generally, the coalescer in the pre-washing step is
beneficial, but not mandatory. It helps to collect and coalesce the
droplets and guides them to the phase interface which typically
results in shorter residence times. Suitable examples of coalescers
include structured or unstructured packings. Structured packings
are for example flat plates, flat vanes, roof-shaped vanes and
vanes with holes in vertical direction. The vanes or plates may be
positioned rectangular or parallel to the main flow direction or
with a slope. Unstructured packings are for example wire mesh,
packings made of rings, spheres, cylinders, irregularly shaped
geometries and weirs like distributor plates that have holes or
slits, vertical plates covering a portion of the main flow path.
The packings can be made of any technically feasible material, e.g.
metals, glass, ceramic, coated metals, lined metals and polymeric
materials like for example PTFE, ETFE, polyethylene (PE),
polyetheretherketone (PEEK), Polypropylene (PP), polyamide (PA) and
polyvinylidenfluoride (PVDF).
[0138] In a preferred embodiment of the invention step pre-a) is
repeated at least once, preferably once.
[0139] A further improved and preferred embodiment of the
pre-washing step is shown in FIG. 10. In step pre-a) of this
double-stage pre-washing step fluid A containing at least one
non-volatile polymer, at least one volatile compound and at least
one hydrophilic compound is fed to the mixing section 30A of a
first separating apparatus 26A, which is equipped with a mixer 32A
and passes through the separating wall 34A into a settling section,
where the mixture separates into an aqueous phase 27A and an
organic phase 28A, whereby the separation is supported by means of
a coalescer 39A. A part of the aqueous phase 27A is removed from
the separating apparatus 26A as fluid C, which is typically
disposed of, with the rest being recycled via the recirculation
line 38A by the action of recirculation pump 36A back into the
mixing section 30A. The organic phase 28A is removed and fed as
fluid B to the mixing section 30B of a second separating apparatus
26B, which is also equipped with a mixer 32B and passes through the
separating wall 34B into a settling section, where the mixture
separates into an aqueous phase 27B and an organic phase 28B,
whereby the separation is supported by means of a coalescer 39B. A
part of the aqueous phase 27B is recycled to the mixing section 30A
of the first separating apparatus 26A as fluid D by the action of
recirculation pump 40 and recirculation line 42, with the rest
being enriched with fresh water E and recycled via the
recirculation line 38B by the action of recirculation pump 36B back
into the mixing section 30B of the second separating apparatus 26B.
The organic phase 28 leaving the second separating apparatus 26B is
subjected to the subsequent process according to steps a) to c) as
fluid F. An advantage of this double-stage pre-washing step is that
fluid F is substantially free of hydrophilic compounds and the
amount of waste water is reduced due to recycling which results in
higher concentration of hydrophilic compounds in fluid C.
[0140] In a preferred embodiment of the invention the separation is
performed at temperatures of more than 40.degree. C. The upper
limit depends on the constitution of the polymer and the
construction of the separating apparatus. Typically the upper limit
is 125.degree. C.
[0141] In a more preferred embodiment of the invention the
separation is performed at temperatures of 40 to 110.degree. C.
preferably at temperatures of 80 to 110.degree..
[0142] Depending on the composition of fluid A and the boiling
points of the components thereof, the separating apparatus may be
designed to be operated under pressure.
[0143] Generally, the efficiency of the pre-washing step increases
with increased temperature.
[0144] In another embodiment of the invention the organic phase 28
leaving the separating apparatus may be pre-heated to facilitate
the free-flow of fluid F. This purpose can also be accomplished by
a heater, whereby heat exchangers as disclosed for heater 2 above
are preferred.
[0145] A further improved and preferred embodiment having
additional heaters for fluid A and fluid F is shown in FIG. 11
which is apart form the heaters identical to FIG. 10. Fluid A is
heated before entering the separating apparatus by heater 25, the
organic phase 28 leaving the second separating apparatus 26B is
heated by heater 44.
[0146] The performance of Step pre-a) is particularly advantageous
for fluids F containing halobutyl rubbers, and in particular for
chlorobutyl and bromobutyl rubbers, since crude halobutyl rubber
solutions often contain high amounts of inorganic halides resulting
from the halogenation of the polymer.
[0147] For example, a fluid A stemming from the bromination of
butyl rubber typically contains inorganic bromide levels of 3,000
to 5,000 ppm calculated on the mass of bromobutyl rubber. Upon
performance of step pre-a) this level can be reduced to less than
500 ppm, preferably to less than 300 ppm and even more preferably
to less than 100 ppm.
[0148] For example, a fluid A stemming front the chlorination of
butyl rubber typically contains inorganic chloride levels of 1,000
to 5,000 ppm calculated on the mass of chlorobutyl rubber. Upon
performance of step pre-a) this level can be reduced to less than
500 ppm, preferably to less than 300 ppm and even more preferably
to less than 100 ppm.
[0149] It was further found that the performance of step pre-a)
allows to significantly reduce the water content of fluid F
compared to fluid A, which contributes to a significantly lower
energy consumption for the subsequent steps a) to c).
[0150] One further embodiment of the invention is shown in FIG. 7.
FIG. 7 shows a basic flow chart and suitable device for the
accomplishment of the process comprising the steps pre-a) and a) to
c).
[0151] In step pre-a) fluid A containing at least one non-volatile
polymer, at least one volatile compound and at least one
hydrophilic compound is fed to the mixing section 30 of the
separating apparatus 26, which is equipped with a mixer 32 and
passes through the separating wall 34 into a settling section,
where the mixture separates into an aqueous phase 27 and an organic
phase 28, whereby the separation is supported by means of a
coalescer 39. A part of the aqueous phase 27 is removed from the
separating apparatus 26 as fluid C, which is typically disposed of,
with the rest being enriched with fresh water E and recycled via
the recirculation line 38 by the action of recirculation pump 36
back into the mixing section 30. The organic phase 28 is removed as
fluid F. In step a) Fluid F is transferred via pump 1 to the heater
2, whereby heated fluid G is obtained. Heated fluid G is fed into
the degassing vessel 4. The vapors emerging from the heated fluid G
are separated and removed by a vacuum line 4.1. After degassing and
separation a concentrated fluid H is obtained, which is removed
from the degassing vessel 4 by means of a pump 4.2.
[0152] In step b), the concentrated fluid H obtained in step a) is
then passed through a reheating unit 6 to obtain a reheated
concentrated fluid L. In step e), the reheated concentrated fluid L
obtained in step b) is passed on to a extruder unit and fed into
the conveying section 16 of the extruder at the feeding point 12.
The conveying section 16 is open to a vent port 15. In the
conveying section 16 a part of the solvent is evaporated and
separated from the reheated concentrated fluid L. The vapors are
removed through the vent port 15 via vapor line 15.1. The conveying
section 16 is terminated by a accumulating section 20. While
passing from the conveying section 16 and the accumulating section
20 to the outlet section 22 the reheated concentrated fluid L
undergoes a transition from the preferably free-flowing reheated
concentrated fluid L to the product P.
[0153] It was further found that the overall energy consumption for
the preparation of halogenated butyl rubbers can be significantly
reduced if fluid A or Fluid F, preferably fluid A is prepared by a
process comprising at least the steps of [0154] I) providing a
reaction medium comprising [0155] a common aliphatic medium
comprising at least 50 wt.-% of one or more aliphatic hydrocarbons
having a boiling point in the range of 45.degree. C. to 80.degree.
C. at a pressure of 1013 hPa, and [0156] a monomer mixture
comprising at least one isoolefin monomer, at least one multiolefin
monomer and either no or at least one other co-polymerizable
monomer [0157] in a mass ratio of monomer mixture to common
aliphatic medium of from 40:60 to 99:1, preferably from 50:50 to
85:15 and even more preferably from 61:39 to 80:20; [0158] II)
polymerizing the monomer mixture within the reaction medium to form
a rubber solution comprising a rubber polymer which is at least
substantially dissolved in the medium comprising the common
aliphatic medium and residual monomers of the monomer mixture;
[0159] III) separating residual monomers of the monomer mixture
from the rubber solution to form a separated rubber solution
comprising the rubber and the common aliphatic medium, [0160] IV)
halogenating the rubber in the separated rubber solution using a
halogenating agent which is in case of a brominating agent
optionally at least partially regenerated by an oxidizing
agent.
[0161] As used herein the term "at least substantially dissolved"
means that at least 70 wt.-%, preferably at least 80 wt-%, more
preferably at least 90 wt.-% and even more preferably at least 95
wt.-% of the rubber polymer obtained according to step II) are
dissolved in the medium.
[0162] In an embodiment of the invention the polymerization
according to step II) and the provision of a solution according to
step I) is effected using a solution polymerization reactor.
Suitable reactors are those known to the skilled in the art and
include commonly known flow-through polymerization reactors.
[0163] Step III) of the process may employ distillation to separate
un-reacted residual monomers, i.e. the isoolefin monomers and the
multiolefin monomers from the medium. This mitigates the formation
of undesirable halogenation byproducts from the unreacted monomers.
The process is conducted at a moderate or relatively high ratio of
monomers to the common aliphatic medium. Typically, the isoolefin
monomers have a significantly lower viscosity than the common
aliphatic medium and therefore, a higher monomer level results in a
lower overall viscosity. Overall energy efficiency and raw material
utilization of the process is improved by eliminating the need to
separate the rubber from a first diluent or solvent used for
polymerization, then re-dissolve it in a second solvent for
bromination and by recycling bromides resulting from bromination
back to a brominating agent. The integrated process according to
the invention therefore provides improved energy and raw material
efficiency and a reduction in the number of process steps as
compared with conventional non-integrated processes for making
halogenated rubbers, in particular bromobutyl rubbers.
[0164] In an embodiment of the invention the bromination according
to step IV) is performed in a continuous process, for example using
a commonly known flow-through halogenation reactor.
[0165] Preferred embodiments of the steps I to IV are exemplarily
described with reference to FIG. 12 which shows a process flow
diagram for a process according to the present invention that
employs purification and optional recycle of un-reacted monomers
following separation thereof from the polymer solution.
[0166] Referring to FIG. 12, a solution polymerization reactor 400
is provided with a feed of monomers FM, comprising isoprene and
isobutylene, and a feed of the common aliphatic medium S via an
optional heat exchanger 100, preferably a recuperative heat
exchanger, and feed cooler 200. The monomers may either be
pre-mixed with the common aliphatic medium or mixed within the
polymerization reactor 400. A catalyst solution, comprising a
carbocationic initiator-activator system of the type used for butyl
rubber polymerizations (e.g. a trivalent metal species, such as
aluminum, and a small amount of water), is pre-mixed with the
common aliphatic medium S in a catalyst preparation unit 300 and
also introduced to the reactor 400. The solution polymerization is
then allowed to occur within the polymerization reactor 400.
Solution polymerization reactors 400 of a type suitable for use in
the present integrated process, along with process control and
operating parameters of such reactors, are described, for example,
in EP 0 053 585 A, which is herein incorporated by reference.
Conversion is allowed to proceed to the desired extent and then a
reaction stopping agent Q, for example water or an alcohol such as
methanol, is added and mixed into the reactor discharge stream
comprising the common aliphatic medium S, un-reacted monomers FM
and butyl rubber IIR in mixer 500. The resulting polymer solution
comprising un-reacted monomers FM i.e. isoprene and isobutylene,
the common aliphatic medium S and butyl rubber IIR is passed
through a recuperative heat exchanger 100 where it is warmed by the
incoming feeds to the reactor, while at the same time helping to
cool these feeds before they enter the final feeds cooler 200. The
warmed polymer solution is then directed to a distillation column
600 for removal of the un-reacted monomers. Once the un-reacted
monomers have been separated as recycling stream FM.sub.R, they
exit from the top of the column 600 and the separated polymer
solution (S, IIR) exits from the bottom of the column 600 to a
solution halogenation reactor 700. Additional common aliphatic
medium S and/or water w may be provided to the bromination reactor
700 in order to provide the desired conditions for bromination. It
is important to note that the same common aliphatic medium used for
polymerization accompanies the butyl rubber through the process to
halogenation and that there is no need to separate the polymer from
the solvent prior to halogenation. A feed of a halogenation agent
HAL and optionally in case the halogenating agent HAL is a
brominating agent optionally an oxidizing agent OX (as described
hereinafter) is also provided to the halogenation reactor 700. The
halogenated butyl rubber (HIIR) exits the reactor in solution (S,
HIIR) and is then finished using finishing equipment 800, as herein
described above and in FIGS. 1 to 11. The common aliphatic medium
removed during the finishing step is sent as recycling stream
S.sub.R to solvent recovery 1100 prior to introduction to solvent
purification section 1200. Additional common aliphatic medium
S.sub.F may be added before purification 1200 or afterwards, if the
medium has already been pre-purified. The purified common aliphatic
medium is recycled back to the recuperative heat exchanger 100 and
final feed cooler 200 for re-use in the process. The un-reacted
monomers separated from the polymer solution in the distillation
column 600 are sent as recycle stream FM.sub.R to monomer recovery
unit 900 and are then purified in monomer purification section 1000
prior to being recycled back to the recuperative heat exchanger 100
and feed cooler 200. Additional fresh monomers M.sub.F may be added
either prior to monomer purification 100 or afterwards, if the
monomers have been pre-purified. The use of a common aliphatic
medium for both polymerization and halogenation reduces
environmental impact and improves economic performance of the
integrated process as compared with conventional approaches.
[0167] The description of the process given hereinabove is
exemplary and can be applied to all common aliphatic media
compositions as well as to all monomer and product compositions
mentioned herein.
[0168] It is within the scope of the present invention that the
composition of the common aliphatic medium may have a slightly
varying composition before and after removal of the un-reacted
monomers due to different boiling points of its components.
[0169] The monomer mixture used to produce the butyl rubber, by
solution polymerization is not limited to a specific isoolefin or a
specific multiolefin or to specific other co-polymerizable
monomers, provided that the individual monomers have boiling points
lower than the aliphatic hydrocarbons of the common aliphatic
medium which are selected from those aliphatic hydrocarbons having
a boiling point in the range of 45.degree. C. to 80.degree. C. at a
pressure of 1013 hPa. It is clear that the boiling point of the
monomers may be higher than 45.degree. C. at a pressure of 1013
hPa, if the aliphatic hydrocarbons of the common aliphatic medium
are selected in such a way that their boiling point is higher than
that of the highest boiling component of the monomer mixture but
still below 80.degree. C. at a pressure of 1013 hPa.
[0170] Preferably, the individual monomers have boiling points
lower than 45.degree. C. at 1013 hPa, preferably lower than
40.degree. C. at 1013 hPa.
[0171] Preferred isoolefins are iso-butene, 2-methyl-1-butene,
3-methyl-1-butene, 2-methyl-2-butene or mixtures thereof. The most
preferred isoolefin is isobutene.
[0172] Preferred multiolefins are isoprene, butadiene or mixtures
thereof. The most preferred multiolefin is isoprene.
[0173] In one embodiment, the monomer mixture may comprise in the
range of from 80.0% to 99.9% by weight, preferably in the range of
from 92.0% to 99.5% by weight of at least one, preferably one
isoolefin monomer and in the range of from 0.1% to 20.0% by weight,
preferably 0.5% to 8.0% by weight of at least one, preferably one
multiolefin monomer. More preferably, the monomer mixture comprises
in the range of from 95.0% to 98.5% by weight of at least one,
preferably one isoolefin monomer and in the range of from 1.5% to
5.0% by weight of at least one, preferably one multiolefin monomer.
Most preferably, the monomer mixture comprises in the range of from
97.0% to 98.5% by weight of at least one, preferably one isoolefin
monomer and in the range of from 1.5% to 3.0% by weight of at least
one, preferably one multiolefin monomer.
[0174] In a preferred embodiment of the invention the ranges given
above apply to monomer mixtures wherein the isoolefin is isobutene
and the multiolefin is isoprene.
[0175] In one embodiment, the multiolefin content of butyl rubbers
produced according to the invention is for example in the range of
0.1 mol % to 20.0 mol %, preferably in the range of 0.5 mol % to
8.0 mol %, more preferably in the range of 1.0 mol % to 5.0 mol %,
yet more preferably in the range of 1.5 mol % to 5 mol % and even
more preferably in the range of 1.8 mol % to 2.2 mol %.
[0176] One of the ways in which the aforementioned viscosity
problems have been overcome is by selecting a high ratio of
monomers to solvent in the polymerization step. Although mass
ratios of up to 60:40 monomers to aliphatic hydrocarbon solvent
have been used in the prior art, in one aspect the present
invention utilizes higher ratios, for example from 61:39 to 80:20,
preferably from 65:35 to 70:30. The presence of higher monomer
levels, which are predominantly C4 compounds and have lower
viscosity than the common aliphatic medium, reduces the solution
viscosity to tolerable limits and also permits a higher solids
level to be achieved during polymerization. Use of higher monomer
levels also allows an acceptable molecular weight to be reached at
a higher temperature than when lower levels of monomer are
employed. The use of higher temperature in turn reduces solution
viscosity and permits greater polymer solids level in the
solution.
[0177] Another one of the ways in which the aforementioned
viscosity problems have been overcome is by selecting the common
aliphatic medium as a solvent. A solvent having a higher content or
consisting of compounds having a boiling point of less than
45.degree. C. or less at 1013 hPa would have a boiling point such
close to the monomers that their separation from the solution would
also result in significant solvent removal.
[0178] The use of a solvent having a higher content or consisting
of compounds having a boiling point of more than 80.degree. C. at
1013 hPa would cause difficulties in the separation from the rubber
after halogenation. The solution viscosity provided by use of such
solvents is also significantly higher than with the common
aliphatic medium, making the solution more difficult to handle and
impeding heat transfer in the reactor, even when provided with the
high monomer to solvent ratios described above.
[0179] In a preferred embodiment of the invention the common
aliphatic medium comprises at least 80 wt.-% of one or more
aliphatic hydrocarbons having a boiling point in the range of
45.degree. C. to 80.degree. C. at a pressure of 1013 hPa,
preferably at least 90 wt.-%, even more preferably at least 95
wt.-% and yet even more preferably at least 97 wt.-%. Aliphatic
hydrocarbons having a boiling point in the range of 45.degree. C.
to 80.degree. C. at a pressure of 1013 hPa include cyclopentane,
2,2-dimethylbutane, 2,3-dimethylbutane, 2-methylpentane,
3-methylpentane, n-hexane, methylcyclopentane and
2,2-dimethylpentane.
[0180] The common aliphatic medium may, for example further
comprise other compounds which are at least substantially inert
under polymerization conditions such as other aliphatic
hydrocarbons like for example heptanes and octanes having a boiling
point of more than 80.degree. C. at a pressure of 1013 hPa,
propanes, butanes, pentanes, cyclohexane as well as
halohydrocarbons such as methylchloride and other chlorinated
aliphatic hydrocarbons which are at least substantially inert under
reaction conditions as well as hydrofluorocarbons whereby
hydrofluorocarbons are for example those represented by the
formula: C.sub.xH.sub.yF.sub.z wherein x is an integer from 1 to
20, alternatively from 1 to preferably from 1 to 3, wherein y and z
are integers and at least one.
[0181] In another preferred embodiment of the invention the common
aliphatic medium is substantially free of halohydrocarbons.
[0182] In another embodiment of the invention the common aliphatic
medium has a content of cyclic aliphatic hydrocarbons of less than
25 wt.-%, preferably less than 20 wt.-%.
[0183] In another embodiment of the invention the common aliphatic
medium has a content of cyclohexane (boiling point: 80.9.degree. C.
at 1013 hPa) of less than 5 wt.-%, preferably less than 2.5
wt.-%.
[0184] As used hereinbefore the term "substantially free of
halohydrocarbons" means a content of halohydrocarbons within the
common aliphatic medium of less than 2 wt.-%, preferably less than
1 wt.-%, more preferably less than 0.1 wt.-% and even more
preferably absence of halohydrocarbons.
[0185] The preferred ratio of monomers to a hydrocarbon solvent is
not calculable in advance, but may be easily determined by very few
routine experiments. Although increasing the amount of monomers
should reduce solution viscosity, making accurate theoretical
predictions of the extent of that reduction is not feasible due in
part to the complex effect on viscosity of the interaction of
various components of the solution at the concentrations and
temperatures employed in the process.
[0186] In one embodiment, the process temperature of step II) is in
the range of -100.degree. C. to -40.degree. C., preferably in the
range of -95.degree. C. to -65.degree. C., more preferably in the
range of -85.degree. C. to -75.degree. C., yet more preferably in
the range of -80.degree. C. to -75.degree. C.
[0187] Although higher temperatures are desirable in that energy
usage for refrigeration and pumping (due to lower viscosity at
higher temperature) are reduced, this generally leads to lower
molecular weight polymers that are not as commercially desirable.
However, due to the use of high monomer to solvent ratios in the
present invention, a reduced but still acceptable molecular weight
can be obtained with higher temperatures.
[0188] Therefore, in an alternative embodiment, temperatures in the
range of -50.degree. C. to lower than -75.degree. C., preferably
-55.degree. C. to -72.degree. C., more preferably -59.degree. C. to
-70.degree. C., yet more preferably -61.degree. C. to -69.degree.
C., are used while still obtaining the desired molecular weight of
butyl rubber.
[0189] The viscosity of the solution at the discharge of reactor 40
is typically and preferably less than 2000 cP, preferably less than
1500 cP, more preferably less than 1000 cP. A most preferred range
of viscosity is from 500 to 1000 cP.
[0190] The solids content of the solution obtained following
polymerization is preferably in the range of from 3 to 25%, more
preferably 10 to 20%, even more preferably from 12 to 18%, yet more
preferably from 14 to 18%, even more preferably from 14.5 to 18%,
still more preferably 15 to 18%, most preferably 16 to 18% by
weight. As described previously, higher solids contents are
preferred, but entail increased solution viscosity. The higher
monomer to solvent ratios used in the present process allow higher
solids contents to be achieved than in the past and advantageously
also permit use of a common aliphatic medium for both
polymerization and bromination.
[0191] As used herein the term "solids content" refers to weight
percent of the polymer obtained according to step II) i.e. in
polymerization and present in the rubber solution.
[0192] In step III), un-reacted residual monomers are removed from
the solution following polymerization preferably using a
distillation process. Distillation processes to separate liquids of
different boiling points are well known in the art and are
described in, for example, the Encyclopedia of Chemical Technology,
Kirk Othmer, 4th Edition, pp. 8-311, which is incorporated herein
by reference.
[0193] The degree of separation is largely dependent upon the
number of trays used in the column. An acceptable and preferred
level of residual monomers in the solution following separation is
less than 20 parts per million by weight. About 40 trays have been
found sufficient to achieve this degree of separation. Separation
of the common aliphatic medium from the monomers is not as critical
and contents of for example up to 10 wt.-% of components of the
common aliphatic medium are acceptable in the overhead stream from
the distillation process. In a preferred embodiment the contents of
components of the common aliphatic medium in the overhead stream
from the distillation process are less than 5 wt.-%, more
preferably less than 1 wt.-%.
[0194] With reference to FIG. 12, the process of the present
invention preferably includes purification of the un-reacted
monomers separated from the polymerization solution using the
distillation column 600. A purification unit 1000 may be provided
for this purpose; alternatively, purification can take place
offsite in a separate purification unit. The purified monomers are
normally recycled back into the process and mixed with fresh
monomers; however, they may alternatively be utilized in a
different process or sold separately. Preferred embodiments of the
process include these optional purification and recycling steps in
order to achieve advantageous overall process economics.
[0195] Purification of monomers may be carried out by passing
through adsorbent columns containing suitable molecular sieves or
alumina based adsorbent materials. In order to minimize
interference with the polymerization reaction, the total
concentration of water and substances such as alcohols and other
organic oxygenates that act as poisons to the reaction are
preferably reduced to less than around 10 parts per million on a
weight basis. The proportion of monomers that are available for
recycle depends on the degree of conversion obtained during the
polymerization process. For example, taking a ratio of monomer to
common aliphatic medium of 66:34, if the solids level in the rubber
solution produced is 10%, then 85% of the monomers are available to
be returned in the recycle stream. If the solids level is increased
to 18%, then 73% of the monomers are available for recycle.
[0196] Following removal of the un-reacted residual monomers, the
butyl polymer is halogenated In step IV). The bromobutyl rubber is
produced using solution phase techniques. The separated rubber
solution comprising the rubber and the common aliphatic medium,
hereinafter also referred to as "cement" is treated with a
halogenating agent and as far as the halogenating agent is a
brominating agent optionally is at least partially regenerated by
an oxidizing agent.
[0197] Supplemental solvent, for example comprising fresh common
aliphatic medium, and/or water may be added to the separated rubber
solution in order to form a cement having the desired properties
for halogenation.
[0198] Halogenation in the common aliphatic medium used during the
polymerization step advantageously saves energy as compared with
the conventional slurry process by eliminating the need for
separating the polymer from the polymerization medium, then
re-dissolving it in a different medium for halogenation.
[0199] Preferably, the amount of halogenating agent is in the range
of from 0.1 to 20%, preferably from 0.1 to 8%, more preferably from
0.5% to 4%, even more preferably from 0.8% to 3%, yet even more
preferably from 1.2 to 2.5%, even still more preferably from about
1.5% to about 2.5%.COPYRGT. and most preferably from 1.5 to 2.5% by
weight of the rubber.
[0200] In another embodiment the quantity of halogenating agent is
0.2 to 1.2 times the molar quantity of double bonds contained in
the rubber, preferably the butyl rubber, preferably 0.3 to 0.8,
more preferably 0.4 to 0.6 times the molar quantity.
[0201] Halogenating agents are for example chlorine and bromination
agents. Suitable bromination agent comprise elemental bromine
(Br.sub.2), interhatogens such as bromine chloride (BrCl) and/or
organo-halide precursors thereto, for example dibromo-dimethyl
hydantoin, N-bromosuccinimide, or the like. The most preferred
brominating agent is molecular bromine (Br.sub.2).
[0202] Where the reaction is conducted with the oxidizing agent
present at the onset of the bromination reaction, hydrogen bromide
may be used as the bromine source. The preferred bromine source is
molecular bromine (Br.sub.2).
[0203] The oxidizing agents which have been found suitable for the
intended purposes of are water soluble materials which contain
oxygen. Preferred oxidizing agents are selected from the group
consisting peroxides and peroxide forming substances as exemplified
by the following substances: hydrogen peroxide, sodium chlorate,
sodium bromate, sodium hypochlorite or bromite, oxygen, oxides of
nitrogen, ozone, urea peroxidate, acids such as pertitanic
perzirconic, perchromic, permolybdic, pertungstic, perboric,
perphosphoric, perpyrophosphoric, persulfates, perchloric,
perchlorate and periodic acids and mixtures of the aforementioned
compounds.
[0204] Such oxidizing agents may either be used in combination with
surfactants or not. In a preferred embodiment no surfactants are
added.
[0205] Suitable surfactants are for example C.sub.6-C.sub.24-alkyl-
or C.sub.6-C.sub.14-aryl-sulfonic acid salts, fatty alcohols and
ethoxylated fatty alcohols and the like materials.
[0206] Preferred oxidizing agents are hydrogen peroxide and
hydrogen peroxide-forming compounds, such as per-acids and sodium
peroxide, whereby hydrogen peroxide is even more preferred.
[0207] For safety reasons, hydrogen peroxide is preferably applied
in form of its aqueous solutions, in particular its aqueous
solutions comprising 25 to 50 wt.-%, preferably 28 to 35 wt.-%,
more preferably around 30 wt.-% of hydrogen peroxide.
[0208] It was found that the lower the water content in the cement
is, the better the bromine utilization and oxidation performance
with hydrogen peroxide is.
[0209] The weight ratio of hydrogen peroxide to water within the
reaction mixture is therefore preferably below 1:100, even more
preferably below 1:50, and yet more preferably below 1:10. In one
embodiment of the invention, the total amount of water present in
the reaction will be provided by the addition of the hydrogen
peroxide solution.
[0210] The amount of oxidizing agent used depends on the amount and
kind of brominating agent used. For example from 0.2 to about 5 mol
of oxidizing agent per mol of brominating agent may be used,
preferably from 0.5 to 3 mol and more preferably from 0.8 to 1.2
mol.
[0211] The oxidizing agent may be introduced into the reaction zone
at the onset of the bromination reaction, it may be added prior to,
concurrently with or subsequent to the addition of the brominating
agent.
[0212] In a preferred embodiment the oxidizing agent is added prior
to the brominating agent to allow its dispersal throughout the
reaction medium the oxidizing agent is added concurrently or before
the brominating agent.
[0213] In another embodiment the oxidizing agent is not added to
the reaction mixture until after at least about 50% of the
brominating agent has been consumed in the bromination
reaction.
[0214] Generally, the halogenation process may be operated at a
temperature of from 0.degree. C. to 90.degree. C., preferably from
20.degree. C. to 80.degree. C. and the reaction time may for
example be from 1 minute to 1 hour, preferably from 1 to 30
minutes. The pressure in the halogenation reactor may be from 0.8
to 10 bar.
[0215] The amount of halogenation during this procedure may be
controlled so that the final polymer has the preferred amounts of
bromine described hereinabove. The specific mode of attaching the
halogen to the polymer is not particularly restricted and those of
skill in the art will recognize that modes other than those
described above may be used while achieving the benefits of the
invention. For additional details and alternative embodiments of
solution phase halogenation processes, see, for example, Ullmann's
Encyclopedia of Industrial Chemistry (Fifth, Completely Revised
Edition, Volume A231 Editors Elvers, et al.) and/or "Rubber
Technology" (Third Edition) by Maurice Morton, Chapter 10 (Van
Nostrand Reinhold Company.COPYRGT. 1987), particularly pp. 297-300,
which are incorporated herein by reference.
[0216] After completion of the halogenation reaction, the polymer
may be recovered by conventional methods, e.g., neutralization with
dilute caustic, water washing and removal of solvent such as by
steam stripping or precipitation using a lower alcohol such as
isopropanol, followed by drying or preferably as disclosed
hereinabove and in FIGS. 8 to 11.
[0217] The invention is in particular advantageous in view of
energy and fresh water consumption. The products obtained are free
of volatile compounds.
EXAMPLES
Analytical Methods
[0218] Water content of fluids F: The sample was put into a
centrifuge and spun for 5 min at 4000 rpm at room temperature. The
water was then collected at the bottom of the vial and weighed.
[0219] Total volatiles concentration: A rubber sample was cut into
small pieces of 2.times.2 mm size. Roughly 30 g of rubber pieces
were put in an alumina crucible. The weight of the crucible and the
rubber was determined. The crucible including the rubber sample was
then placed in a vacuum oven at a vacuum level of 130 hPa for 60
min at a temperature of 105.degree. C. After drying, the crucible
was placed in an exsiccator and let cool down for 30 min. The
crucible was then weighed again. The loss in weight was
determined.
[0220] Residual solvent concentration in product P: The residual
solvent concentration in the product was determined by headspace
gas chromatography. A weighed portion (0.5+-0.005 g) of sample was
placed in a headspace vial, and a measured amount of solvent (1,2
dichlorobenzene, ODCB) was added. The vial was sealed and shaken
until the rubber was dissolved. The vial was heated until the
volatile organic compounds were distributed at equilibrium between
the sample and the gas phase in the vial (headspace). An aliquot of
the headspace gas was injected into a stream of carrier gas, which
carries the sample along a chromatographic column. Standards of
known composition were used to calibrate the GC. Toluene was added
to the solvent for use as an Internal Standard.
[0221] Residual water concentration in product P: The total
volatiles concentration is the sum of water, solvents and monomers.
As the monomer concentration is usually less then 0.0005 wt %, the
water content can be determined by subtracting the solvent
concentration from the total volatiles concentration.
[0222] Solvent concentration in fluids: The concentration of
solvents in fluids were measured using gas chromatography. The
internal standard was isooctane. The sample was diluted with
toluene and then injected into the gas chromatograph. The gas
chromatography was performed on a HP 6890 chromatograph, with
following specifications: [0223] column type DB-5 of J&W,
length 60 m, diameter 0.23 mm, film thickness 1.0 .mu.m [0224]
injector temp.: 250.degree. C. [0225] detector temp.: 350.degree.
C. [0226] carrier gas: Helium [0227] column pressure: 96 kPa [0228]
detector: FID
[0229] Viscosity of fluids: The viscosity was measured in a
rotational rheometer of cone-plate type. All given viscosities
refer to the extrapolated zero shear viscosity.
Example 1
Polymerization and Distillation
[0230] Key elements of the process described in FIG. 13 have been
operated at pilot scale with reactors of 20 litre total capacity
running in a continuous mode. Feeds to the reactors were 38.7 kg/h
of isobutene, 0.9 kg/h of isoprene and 20.0 kg/h of hexane giving a
monomer/hexane mass ratio of 66:34. The reaction temperature used
was -65.degree. C. and a solution having a solids content of 16 wt
% was produced. This material had a weight average molecular weight
of about 440 kg/mol and an isoprene content of about 1.7 mol-%. The
solution from the reactors was fed to a distillation column with 40
trays and separation of the monomers from the rubber solution was
performed. The solution was preheated to 42.degree. C. and a
re-boiler was used at the bottom of the column to maintain a bottom
temperature of 113.degree. C. A reflux condenser was used to return
part of the overhead stream to the top of the column maintaining a
temperature there of 36.degree. C. The separation achieved in the
column left less than 10 ppm of residual isoprene monomer in the
separated rubber solution and 1.2% of hexane in the overhead
monomer stream. The separated monomers were purified, then
re-introduced to the solution polymerization reactor. The separated
rubber solution in the hexane solvent was such that bromination
could be accomplished by conventional means with addition of
supplemental hexane solvent.
Example 2
Halogenation
[0231] The separated rubber solution of Example 1 was halogenated
using a continuous pilot scale bromination equipment. Supplemental
solvent in an amount of 10% was added to the separated rubber
solution in order to lower the viscosity. To simulate plant
conditions, supplemental water was added to the solution and
allowed to disperse throughout the reaction medium. 30 wt.-%
hydrogen peroxide in water (at a molar ratio of 1:1 with bromine to
be added) was introduced into this solution and the resulting
mixture was agitated at 50.degree. C. for up to 2 minutes prior to
the addition of bromine. The amount of bromine added was 24 kg per
ton of base rubber (=65% of standard, non-recovery bromination
amount). After a reaction period of up to 30 minutes, caustic
solution was added to the reaction mixture to neutralize any
residual hydrogen bromide, bromine and hydrogen peroxide. The
neutralized cement was used for example 3 as fluid (A).
Example 3
Pre-Washing
[0232] The crude bromobutyl rubber solution, hereinafter denoted as
fluid (A) contained two phases: an aqueous phase (56 wt %) and an
organic phase (44 wt %). The overall ratio of bromobutyl rubber
with respect to hexanes in the organic phase alone was constant
throughout the examples, being 22 wt % bromobutyl rubber and about
78 wt % hexanes. The bromobutyl rubber, contained in fluid (A) had
the following properties, once finished and dried: Mooney (ML 1+8,
125.degree. C.) of 32.+-.4, bound bromine content 1.8.+-.0.2 wt
%.
[0233] Fluid (A) further comprised certain additives, the
concentration being given as mass fraction with respect to the
rubber mass (phr=parts per hundred parts of rubber):
[0234] ESBO: 1 to 1.6 phr, calcium stearate 1.3 to 1.7 phr, Irganox
0.03 to 0.1 phr
[0235] The aqueous phase had a typical pH-value of 9.5. In addition
to the additives, fluid (A) comprised inorganic components like
bromides, chlorides, calcium, sodium, aluminum and small amounts of
other inorganic components.
[0236] The experiment was carried out using a glass vessel having a
volume of 11 and performed batchwise. The vessel was equipped with
a stirrer.
[0237] The water content in the organic phase was determined as
described above.
[0238] A sample of fluid (A) was taken and left settling. The
aqueous phase and the organic phase were analyzed. The aqueous
phase contained 4940 mg/l of inorganic bromides. The organic phase
contained 20 wt % bromobutyl rubber, 68 wt % hexane and 12 wt %
water. The total inorganic bromine concentration in the organic
phase was 0.15 wt % (1500 ppm).
Examples 4 to 23
Concentration and Direct Evaporation
[0239] The fluid (F) containing butyl rubber used as a feedstock
for examples 4 to 23 was obtained from two different sources:
Preparation of Fluid F1
[0240] A crude butyl rubber solution was taken from a commercial
production plant, allowed to settle several hours and the organic
phase separated from the bulk aqueous phase. The organic phase was
then used to perform the experiments as fluid (F1). Fluid (F1)
contained 20 wt % rubber, 70 wt % hexanes and 10 wt % water
calculated on 100 wt % of these three components and was thus very
similar to the organic phase obtained in example 3. The
concentration of additives with respect to the bromobutyl rubber
fraction was:
[0241] ESBO: 1 to 1.6 phr, Calcium stearate: 1.3 to 1.7 phr and
Irganox: 0.03 to 0.1 phr
[0242] The bromobutyl rubber, dissolved in the fluid (F1), had the
following properties, once finished and dried: Mooney (ML 1+8,
125.degree. C.) of 28 to 36, Bound bromine content of 1.6 to 2.0 wt
%.
[0243] The viscosity of Fluid F1 at 60.degree. C. was 1,760
mPa*s
Preparation of Fluid F2
[0244] Commercially available bromobutyl rubber with a Mooney (ML
1+8, 125.degree. C.) of 28 to 36, a bromine content of 1.6 to 2.0
wt % and an organic volatile concentration of <0.7 wt % was
dissolved in technical hexane whereby a fluid (F2) was obtained
containing 20 wt % rubber, 79 wt % hexanes and 1 wt % water
calculated on 100 wt % of these three components. Therefore also
fluid F2 was very similar to the organic phase obtained in example
3. The concentration of additives with respect to the bromobutyl
rubber fraction was:
[0245] ESBO: 1 to 1.6 phr, Calcium stearate 1.3 to 1.7 phr and
Irganox: 0.03 to 0.1 phr
Examples 4 to 8
Concentration
The Concentrator Unit
[0246] The concentrator unit used for the examples was similar to
the one shown in FIG. 1. A piston pump was used to pump the fluid
F1, which was prepared as described above, to heater (2). The
heater (2) was a single tube-in-tube type heat exchanger. The
internal pipe was equipped with a static mixer of Kenics type, the
diameter of the internal pipe was 15 mm. The tube was heated by a
tube shaped shell. The heating medium was heating oil (Marlotherm).
A pressure relief valve (3) was installed prior to the degassing
vessel (4), the pressure upstream of the valve was controlled
automatically to a set point value. This set point was chosen so
that boiling in the heated fluid (G1) was prevented. The heated
fluid (G) was introduced into the degassing vessel (4) from the
top. The conical outlet of the degassing vessel (4) was equipped
with a pump (4.2), which was a combination of an extruder type pump
and a gear pump. This combination had the advantage of being able
to handle high viscosities and to build up high pressures. Samples
were taken from the concentrated fluid (H) to investigate the
concentration and viscosity after the concentration stage.
Example 4
[0247] The heating medium of the heater 2 was set to 125.degree.
C., the pressure in the separating vessel 4 was atmospheric (1013
hPa). The concentrated fluid H in the bottom of the separating
vessel 4 was a free flowing foamy liquid, as observed through a
sight glass and could be easily conveyed from the separating vessel
using the extraction pump 4.2 as described above. The concentrated
fluid H had a hexane concentration of 71 wt % and a viscosity of
4,840 mPa*s measured at 60.degree. C.
Example 5
[0248] The feedstock, fluid F1, and the concentration unit were the
same as in example 8. The heating medium of the heater 2 was set to
155.degree. C., the pressure in the separating vessel 4 was
atmospheric (1013 hPa). The concentrated fluid H in the bottom of
the separating vessel 4 was a free flowing foamy liquid, as
observed through a sight glass and could be easily conveyed from
the separating vessel using the extraction pump 4.2 as described
above. The concentrated fluid H had a hexane concentration of 53 wt
% and a viscosity of 65,000 mPa*s measured at 60.degree. C.
Example 6
[0249] The feedstock, fluid F1, and the concentration unit were the
same as in example 8. The heating medium of the heater 2 was set to
170.degree. C., the pressure in the separating vessel 4 was
atmospheric (1013 hPa). The concentrated fluid H in the bottom of
the separating vessel 4 was a free flowing foamy liquid, as
observed through a sight glass and could be conveyed from the
separating vessel without plugging or product buildup using the
extraction pump 4.2 as described above. The concentrated fluid H
had a hexane concentration of 42 wt % and a viscosity of 317,700
mPa*s measured at 60.degree. C.
Example 7
[0250] The feedstock, fluid F1, and the concentration unit were the
same as in example 8. The heating medium of the heater 2 was set to
170.degree. C., the pressure in the separating vessel 4 was 500
hPa. The concentrated fluid H in the bottom of the separating
vessel 4 was a free flowing foamy liquid, as observed through a
sight glass and could be conveyed from the separating vessel using
the extraction pump 4.2 as described above. Only little product
buildup was observed in the conical outlet section of the
separating vessel 4. The concentrated fluid H had a hexane
concentration of 20 wt % and a viscosity of 7,600,000 mPa*s
measured at 60.degree. C.
Example 8
[0251] The feedstock, fluid F1, and the concentration unit were the
same as in example 8. The heating medium of the heater 2 was set to
170.degree. C., the pressure in the separating vessel 4 was 230
hPa. The concentrated fluid H in the bottom of the separating
vessel 4 was a free flowing foamy liquid, as observed through a
sight glass and could be conveyed from the separating vessel using
the extraction pump 4.2 as described above. Some product buildup
was observed in the conical outlet section of the separating vessel
4. The concentrated fluid H had a hexane concentration of 15 wt %
and a viscosity of 15,600,000 mPa*s measured at 60.degree. C.
[0252] The results of examples 4 to 8 showing the performance of
the concentration stage are summarized in table 1.
TABLE-US-00001 TABLE 1 T [.degree. C.] P in [hPa] Hexane Viscosity
[mPa*s] at in degassing content [wt %] at 60.degree. C. Example
heater 2* vessel 4 of fluid H* of fluid H 4 125.degree. C. 1013 71%
4,840 5 155.degree. C. 1013 53% 65,000 6 170.degree. C. 1013 42%
317,700 7 170.degree. C. 500 20% 7,600,000 8 170.degree. C. 230 15%
15,600,000 *temperature set for the heating medium
Examples 9 to 15
Concentration and Extrusion
The Device
[0253] The device used for the examples was similar to the one
shown in FIG. 5. A piston pump was used to pump the fluid F to
heater 2. The heater 2 was a single tube-in-tube type heat
exchanger. The internal pipe was equipped with a static mixer of
Kenics type, the diameter of the internal pipe was 15 mm. The tube
was heated by a tube shaped shell. The heating medium was heating
oil (Marlotherm). A pressure relief valve 3 was installed prior to
the degassing vessel 4, the pressure upstream of the valve was
controlled automatically to a set point value. This set point was
chosen so that boiling in the heated fluid G was prevented. The
heated fluid G was introduced into the degassing vessel 4 from the
top. The conical outlet of the degassing vessel 4 was equipped with
a pump 4.2, which was a combination of an extruder type pump and a
gear pump.
[0254] In step b), the concentrated fluid H obtained in step a) was
then passed through a reheating unit 6 which was a single
tube-in-tube type heat exchanger. The internal pipe diameter was 20
mm, the internal pipe was equipped with a static mixer of type SMX.
Heating was accomplished by a tube shell using a heating oil
(Marlotherm) as heating medium.
[0255] In step c) the reheated concentrated fluid L was fed into
the extruder unit. The extruder of the extruder unit was a
co-rotating twin screw extruder with a screw diameter of 32 mm and
a screw length of 1260 mm. The extruder unit further comprised a
nozzle as a pressure control device 7, (see FIG. 7) upstream the
feeding point 12 of the extruder, three extruder degassing
sections, whereby the feeding point 12 was located at the first
extruder degassing section, whereby the first extruder degassing
section comprised a conveying section 16A, a rear vent port 13
connected to a vapor line 13.1 in upstream direction and whereby
the extruder unit further comprised two downstream extruder
degassing sections each comprising a conveying section 16B and 16
C, a vent port 15A and 15B, whereby the vent ports 15A and 15B were
each connected to a vapour line 15.1A and 15.1B and whereby each of
the conveying sections 16A, 16B and 16C was terminated by a
accumulating section 18A, 18B and 20 and whereby the extruder unit
further comprised an outlet section 22.
[0256] Each of the sections, in particular the conveying sections
could be independently heated through the barrel of the extruder in
order to control the temperature of the rubber anywhere in the
extruder.
[0257] The rear vent port 13 was connected to a condenser via a
first vapor line 13.1. The condenser was a plate type heat
exchanger and further connected to a liquid ring vacuum pump. The
other vapor lines 15.1A and 15.1B were connected to a condensing
system comprising a screw type dry running vacuum pump.
[0258] The first accumulating section 18A was made of kneading
blocks, the second accumulating section 18B was made of kneading
blocks and a back conveying element. Both accumulating sections 18A
and 18B were designed to allow the injection of a stripping
agent.
[0259] A sight glass was installed in the vent port 15.1B to allow
the observation of the conveying behavior and of the product
properties in the conveying section 16C.
[0260] The kneading zone 20 and outlet section 22 were combined
into one functional section. The accumulating section zone was
composed of a die plate and a nozzle forming a strand of rubber
which was formed into rubber crumbs at the outlet section.
Example 9
[0261] Fluid F2 as described above was used as feedstock (fluid F).
The throughput of fluid F2 was set to 20 kg/h, which corresponds to
4.4 kg/h of the final bromobutyl rubber product. The heating
temperature of the heater 2 was set to 154.degree. C., the pressure
in the separating vessel 4 to 626 hPa. The temperature of the
heating medium of the reheating unit 6 was set to 153.degree. C.,
the pressure in the rear vent port 13 was 626 hPa. The barrel
temperature of the extruder was 81.degree. C. The pressure in the
second and third vent port 15A and 15B was lowered to 6 hPa.
Nitrogen was fed into the accumulating section 18B as a stripping
agent at a rate of 0.85 wt % with respect to mass of the final
bromobutyl rubber product. Through the sight glass in the
separating vessel 4 it was observed, that the concentrated fluid H
was still a free flowing fluid. Through the sight glass at the last
vent port of the extruder 15B it could be observed that the rubber
had already changed to a crumbly state in the extruder. The crumbs
appeared white and were permanently drawn in and kneaded by the
action of the screw shafts. At the outlet section 22 a strand of
rubber was produced, which was then cut into crumbs or chunks of
rubber. The final bromobutyl rubber product P collected at the
outlet section was analyzed to determine the hexane and total
volatiles concentration.
Total volatiles: 0.89 wt %
Hexane: 0.65 wt %
Water: 0.24 wt %.
Example 10
[0262] Fluid F2 as described above was used as feedstock (fluid F).
The throughput of fluid F2 was set to 5 kg/h, which corresponds to
1.1 kg/h of the final bromobutyl rubber product. The heating
temperature of the heater 2 was set to 157.degree. C., the pressure
in the separating vessel 4 to 633 hPa. The temperature of the
heating medium of the reheating unit 6 was set to 156.degree. C.,
the pressure in the rear vent port 13 was 633 hPa. The barrel
temperature of the extruder was 81.degree. C. The pressure in the
second and third vent port 15A and 15B was lowered to 6 hPa.
Nitrogen was fed into the accumulating section 18B as a stripping
agent at a rate of 3.41 wt % with respect to mass of the final
bromobutyl rubber product. Through the sight glass in the
separating vessel 4 it was observed, that the concentrated fluid H
was still a free flowing fluid. Through the sight glass at the last
vent port of the extruder 15B it could be observed that the rubber
had already changed to a crumbly state in the extruder. The crumbs
appeared white and were permanently drawn in and kneaded by the
action of the screw shafts. At the outlet section 22 a strand of
rubber was produced, which was then cut into crumbs or chunks of
rubber. The final bromobutyl rubber product P collected at the
outlet section was analyzed to determine the hexane and total
volatiles concentration.
Total volatiles: 0.72 wt %
Hexane: 0.56 wt %
Water: 0.16 wt %.
Example 11
[0263] Fluid F2 as described above was used as feedstock (fluid F).
The throughput of fluid F2 was set to 10 kg/h, which corresponds to
2.2 kg/h of the final bromobutyl rubber product. The heating
temperature of the heater 2 was set to 156.degree. C., the pressure
in the separating vessel 4 to 318 hPa. The temperature of the
heating medium of the reheating unit 6 was set to 156.degree. C.,
the pressure in the rear vent port 13 was 318 hPa. The barrel
temperature of the extruder was 81.degree. C. The pressure in the
second and third vent port 15A and 15B was lowered to 12 hPa.
Nitrogen was fed into the accumulating section 18B as a stripping
agent at a rate of 1.70 wt % with respect to mass of the final
bromobutyl rubber product. Through the sight glass in the
separating vessel 4 it was observed, that the concentrated fluid H
was still a free flowing fluid. Through the sight glass at the last
vent port of the extruder 15B it could be observed that the rubber
had already changed to a crumbly state in the extruder. The crumbs
appeared white and were permanently drawn in and kneaded by the
action of the screw shafts. At the outlet section 22 a strand of
rubber was produced, which was then cut into crumbs or chunks of
rubber. The final bromobutyl rubber product P collected at the
outlet section was analyzed to determine the hexane and total
volatiles concentration.
Total volatiles: 0.80 wt %
Hexane: 0.40 wt %
Water: 0.40 wt %.
Example 12
[0264] Fluid F2 as described above was used as feedstock (fluid F).
The throughput of fluid F2 was set to 10 kg/h, which corresponds to
2.2 kg/h of the final bromobutyl rubber product. The heating
temperature of the heater 2 was set to 155.degree. C., the pressure
in the separating vessel 4 to 475 hPa. The temperature of the
heating medium of the reheating unit 6 was set to 156.degree. C.,
the pressure in the rear vent port 13 was 475 hPa. The barrel
temperature of the extruder was 100.degree. C. The pressure in the
second and third vent port 15A and 15B was lowered to 11 hPa. No
stripping agent was fed into the accumulating section 18B. Through
the sight glass in the separating vessel 4 it was observed, that
the concentrated fluid H was still a free flowing fluid. Through
the sight glass at the last vent port of the extruder 15B it could
be observed that the rubber had already changed to a crumbly state
in the extruder. The crumbs appeared white and were permanently
drawn in and kneaded by the action of the screw shafts. At the
outlet section 22 a strand of rubber was produced, which was then
cut into crumbs or chunks of rubber. The final bromobutyl rubber
product P collected at the outlet section was analyzed to determine
the hexane and total volatiles concentration.
Total volatiles: 0.97 wt %
Hexane: 0.58 wt %
Water: 0.39 wt %.
Example 13
[0265] Fluid F2 as described above was used as feedstock (fluid F).
The throughput of fluid F2 was set to 10 kg/h, which corresponds to
2.2 kg/h of the final bromobutyl rubber product. The heating
temperature of the heater 2 was set to 155.degree. C., the pressure
in the separating vessel 4 to 475 hPa. The temperature of the
heating medium of the reheating unit 6 was set to 156.degree. C.,
the pressure in the rear vent port 13 was 475 hPa. The barrel
temperature of the extruder was 100.degree. C. The pressure in the
second and third vent port 15A and 15B was lowered to 11 hPa. Water
was fed into the accumulating section 18B as a stripping agent at a
rate of 4.09 wt % with respect to mass of the final bromobutyl
rubber product. Through the sight glass in the separating vessel 4
it was observed, that the concentrated fluid H was still a free
flowing fluid. Through the sight glass at the last vent port of the
extruder 15B it could be observed that the rubber had already
changed to a crumbly state in the extruder. The crumbs appeared
white and were permanently drawn in and kneaded by the action of
the screw shafts. At the outlet section 22 a strand of rubber was
produced, which was then cut into crumbs or chunks of rubber. The
final bromobutyl rubber product P collected at the outlet section
was analyzed to determine the hexane and total volatiles
concentration.
Total volatiles: 0.45 wt %
Hexane: 0.31 wt %
Water: 0.14 wt %.
Example 14
[0266] Fluid F2 as described above was used as feedstock (fluid F).
The throughput of fluid F2 was set to 10 kg/h, which corresponds to
2.2 kg/h of the final bromobutyl rubber product. The heating
temperature of the heater 2 was set to 155.degree. C., the pressure
in the separating vessel 4 to 475 hPa. The temperature of the
heating medium of the reheating unit 6 was set to 156.degree. C.,
the pressure in the rear vent port 13 was 475 hPa. The barrel
temperature of the extruder was 130.degree. C. The pressure in the
second and third vent port 15A and 15B was lowered to 11 hPa. Water
was fed into the accumulating section 18B as a stripping agent at a
rate of 4.09 wt % with respect to mass of the final bromobutyl
rubber product. Through the sight glass in the separating vessel 4
it was observed, that the concentrated fluid H was still a free
flowing fluid. Through the sight glass at the last vent port of the
extruder 15B it could be observed that the rubber had already
changed to a crumbly state in the extruder. The crumbs appeared
white and were permanently drawn in and kneaded by the action of
the screw shafts. At the outlet section 22 a strand of rubber was
produced, which was then cut into crumbs or chunks of rubber. The
final bromobutyl rubber product P collected at the outlet section
was analyzed to determine the hexane and total volatiles
concentration.
Total volatiles: 0.22 wt %
Hexane: 0.13 wt %
Water: 0.09 wt %.
Example 15
[0267] Fluid F2 as described above was used as feedstock (fluid F).
The throughput of fluid F2 was set to 10 kg/h, which corresponds to
2.2 kWh of the final bromobutyl rubber product. The heating
temperature of the heater 2 was set to 155.degree. C., the pressure
in the separating vessel 4 to 475 hPa. The temperature of the
heating medium of the reheating unit 6 was set to 156.degree. C.,
the pressure in the rear vent port 13 was 475 hPa. The barrel
temperature of the extruder was 160.degree. C. The pressure in the
second and third vent port 15A and 15B was lowered to 11 hPa. Water
was fed into the accumulating section 18B as a stripping agent at a
rate of 4.09 wt % with respect to mass of the final bromobutyl
rubber product. Through the sight glass in the separating vessel 4
it was observed, that the concentrated fluid H was still a free
flowing fluid. Through the sight glass at the last vent port of the
extruder 15B it could be observed that the rubber had already
changed to a crumbly state in the extruder. The crumbs appeared
white and were permanently drawn in and kneaded by the action of
the screw shafts. At the outlet section 22 a strand of rubber was
produced, which was then cut into crumbs or chunks of rubber. The
final bromobutyl rubber product P collected at the outlet section
was analyzed to determine the hexane and total volatiles
concentration.
Total volatiles: 0.09 wt %
Hexane: 0.04 wt %
Water: 0.05 wt %.
[0268] The results of examples 9 to 15 are summarized in tables
2a), b) and c).
TABLE-US-00002 TABLE 2a Concentration Unit Process conditions in
the Concentration and Reheating Units Throughput [kg/h] Throughput*
P [hPa] in T [.degree. C.] at of fluid [kg/h] of T [.degree. C.] at
degassing reheating Example F2 Product P heater 2* vessel 4 unit 6*
9 20 4.4 154 626 153 10 5 1.1 157 633 156 11 10 1.1 156 318 156 12
10 1.1 155 475 155 13 10 1.1 155 475 155 14 10 1.1 155 475 155 15
10 1.1 155 475 155 * temperature set for the heating medium
TABLE-US-00003 TABLE 2b Extruder Unit Process conditions in the
Extruder Unit P [hPa] at P [hPa] at P [hPa] at Stripping agent vent
port vent port vent port at section Amount Example 13 15A 15B 18B
[wt %]* 9 626 6 6 nitrogen 0.85 10 633 6 6 nitrogen 3.41 11 318 12
12 nitrogen 1.70 12 475 11 11 none -- 13 475 11 11 water 4.09 14
475 11 11 water 4.09 15 475 11 11 water 4.09 *wt % with respect to
bromobutyl rubber product
TABLE-US-00004 TABLE 2c Results Contents of volatiles in the final
product Hexane Water Total Volatiles Example [wt %] [wt %]* [wt %]
9 0.65 0.24 0.89 10 0.56 0.16 0.72 11 0.40 0.40 0.80 12 0.58 0.39
0.97 13 0.31 0.14 0.45 14 0.13 0.09 0.22 15 0.04 0.05 0.09
*Difference of Total Volatiles and Hexane content
Examples 16 to 19
Concentration and Extrusion
The Device
[0269] The device used for the examples was similar to the one
shown in FIG. 6 and identical to the one described for examples 13
to 19, except that: [0270] The extruder unit comprised a fourth
degassing zone, comprising a fourth conveying section 16D and a
fourth vent port 15C equipped with a vapor line 15.1C which was
connected to a previous vapor line 15.1B. [0271] The sight glass to
observe the product behavior was part of the vent port 15C instead
of the vent port 15B. [0272] The third accumulating section 18C was
made of kneading elements and a back conveying element similar to
the second accumulating section 18B and was also designed to allow
the injection of a stripping agent. [0273] The final kneading zone
20 comprised kneading and back conveying elements [0274] The outlet
section 22 just comprised screw conveying elements and an open
outlet.
General Procedure
[0275] Fluid F2 as described above was used as feedstock (fluid F).
The throughput of fluid F2 was set to 10 kWh, which corresponds to
2.2 kWh of the final bromobutyl rubber product. The heating
temperature of the heater 2 was set to 160.degree. C., the pressure
in the separating vessel 4 to 450 hPa. The temperature of the
heating medium of the reheating unit 6 was set to 160.degree. C.,
the pressure in the rear vent port 13 was 450 hPa. The barrel
temperature of the extruder was 130.degree. C. The pressure in the
second, third and fourth vent port 15A, 15B and 15C was lowered to
6 hPa.
[0276] The types and amounts of stripping agents injected at the
accumulating sections 18B and 18C were varied as can be seen in
table 3b). Through the sight glass in the separating vessel 4 it
was observed for each experiment, that the concentrated fluid H was
still a free flowing fluid. Through the sight glass at the last
vent port of the extruder 15C it could be observed that the rubber
had already changed to a crumbly state in the extruder. The crumbs
appeared white and were permanently drawn in and kneaded by the
action of the screw shafts. At the outlet section 22 rubber crumbs
of a size of roughly 2-6 mm were formed. The final bromobutyl
rubber product P collected at the outlet section was analyzed to
determine the hexane and total volatiles concentration.
[0277] The process conditions and results are given in tables 3a),
b) and c).
TABLE-US-00005 TABLE 3a Concentration Unit Process conditions in
the Concentration and Reheating Units Throughput Throughput* P
[hPa] T [.degree. C.] [kg/h] [kg/h] T [.degree. C.] in at of fluid
of Product at degassing reheating Example F2 P heater 2* vessel 4
unit 6* 16 to 19 10 1.1 160 450 160 *temperature set for the
heating medium
TABLE-US-00006 TABLE 3B Extruder Unit Process conditions in the
Extruder Unit* Strip- Strip- P [hPa] P [hPa] ping ping at at vent
agent agent vent ports at at port 15A, 15B section Amount section
Amount Example 13 and 15C 18B [wt %]** 18B [wt %]** 19 450 6 none
-- none -- 20 450 6 nitrogen 1.70 none -- 21 450 6 none -- water
2.73 22 450 6 nitrogen 1.70 water 2.73 *The barrel temperature of
the extruder was set to 130.degree. C. in all examples **wt % with
respect to bromobutyl rubber product
TABLE-US-00007 TABLE 3c Results Contents of volatiles in the final
product Hexane Water Total Volatiles Example [wt %] [wt %] [wt %]
19 0.03 0.08 0.11 20 0.02 0.08 0.10 21 0.03 0.12 0.15 22 0.02 0.07
0.09
Examples 20 to 23
Concentration and Extrusion
The Device
[0278] The device used for the examples was identical to the one
described for examples 16 to 19
Example 20
[0279] Fluid F1 as described above was used as feedstock (fluid F).
The throughput of fluid F1 was set to 20 kg/h, which corresponds to
4.4 kg/h of the final bromobutyl rubber product. The heating
temperature of the heater 2 was set to 139.degree. C., the pressure
in the separating vessel 4 to 756 hPa. The temperature of the
heating medium of the reheating unit 6 was set to 153.degree. C.,
the pressure in the rear vent port 13 was 147 hPa. The barrel
temperature of the extruder was 130.degree. C. The pressure in the
second vent port 15A was lowered to 270 hPa, the pressure in the
third and fourth vent port 15b and 15C was lowered to 40 hPa. Water
was fed into each of the accumulating sections 18B and 18C as a
stripping agent at a rate of 1.36 wt % with respect to the mass of
the final bromobutyl rubber product. Through the sight glass in the
separating vessel 4 it was observed, that the concentrated fluid H
was still a free flowing fluid. Through the sight glass at the last
vent port of the extruder 15C it could be observed that the rubber
had already changed to a crumbly state in the extruder. The crumbs
appeared white and were permanently drawn in and kneaded by the
action of the screw shafts. At the outlet section 22 rubber crumbs
of a size of roughly 2-6 mm were formed. The final bromobutyl
rubber product P collected at the outlet section was analyzed to
determine the hexane and total volatiles concentration.
Total volatiles: 0.200 wt %
Hexane: 0.080 wt %
Water: 0.120 wt %.
Example 21
[0280] Fluid F1 as described above was used as feedstock (fluid F).
The throughput of fluid F1 was set to 20 kWh, which corresponds to
4.4 kg/h of the final bromobutyl rubber product. The heating
temperature of the heater 2 was set to 157.degree. C., the pressure
in the separating vessel 4 to 869 hPa. The temperature of the
heating medium of the reheating unit 6 was set to 147.degree. C.,
the pressure in the rear vent port 13 was 869 hPa. The barrel
temperature of the extruder was 130.degree. C. The pressure in the
second vent port 15A was lowered to 270 hPa, the pressure in the
third and fourth vent port 15b and 15C was lowered to 40 hPa. Water
was fed into each of the accumulating sections 18B and 18C as a
stripping agent at a rate of 2.73 wt % with respect to the mass of
the final bromobutyl rubber product. Through the sight glass in the
separating vessel 4 it was observed, that the concentrated fluid H
was still a free flowing fluid. Through the sight glass at the last
vent port of the extruder 15C it could be observed that the rubber
had already changed to a crumbly state in the extruder. The crumbs
appeared white and were permanently drawn in and kneaded by the
action of the screw shafts. At the outlet section 22 rubber crumbs
of a size of roughly 2-6 mm were formed. The final bromobutyl
rubber product P collected at the outlet section was analyzed to
determine the hexane and total volatiles concentration.
Total volatiles: 0.260 wt %
Hexane: 0.092 wt %
Water: 0.168 wt %.
Example 22
[0281] Fluid F1 as described above was used as feedstock (fluid F).
The throughput of fluid F1 was set to 20 kg/h, which corresponds to
4.4 kg/h of the final bromobutyl rubber product. The heating
temperature of the heater 2 was set to 157.degree. C., the pressure
in the separating vessel 4 to 796 hPa. The temperature of the
heating medium of the reheating unit 6 was set to 147.degree. C.,
the pressure in the rear vent port 13 was 796 hPa. The barrel
temperature of the extruder was 130.degree. C. The pressure in the
second vent port 15A was lowered to 140 hPa, the pressure in the
third and fourth vent port 15b and 15C was lowered to 40 hPa. Water
was fed into each of the accumulating sections 18B and 18C as a
stripping agent at a rate of 1.29 wt % with respect to the mass of
the final bromobutyl rubber product. Through the sight glass in the
separating vessel 4 it was observed, that the concentrated fluid H
was still a free flowing fluid. Through the sight glass at the last
vent port of the extruder 15C it could be observed that the rubber
had already changed to a crumbly state in the extruder. The crumbs
appeared white and were permanently drawn in and kneaded by the
action of the screw shafts. At the outlet section 22 rubber crumbs
of a size of roughly 2-6 mm were formed. The final bromobutyl
rubber product P collected at the outlet section was analyzed to
determine the hexane and total volatiles concentration.
Total volatiles: 0.180 wt %
Hexane: 0.099 wt %
Water: 0.081 wt %.
Example 23
[0282] Fluid F1 as described above was used as feedstock (fluid F).
The throughput of fluid F1 was set to 20 kg/h, which corresponds to
4.4 kg/h of the final bromobutyl rubber product. The heating
temperature of the heater 2 was set to 157.degree. C., the pressure
in the separating vessel 4 to 791 hPa. The temperature of the
heating medium of the reheating unit 6 was set to 147.degree. C.,
the pressure in the rear vent port 13 was 791 hPa. The barrel
temperature of the extruder was 130.degree. C. The pressure in the
second vent port 15A was lowered to 140 hPa, the pressure in the
third and fourth vent port 15b and 15C was lowered to 40 hPa.
Nitrogen was fed into the first accumulating section 18B at a rate
of 0.89 wt % with respect to the mass of the final bromobutyl
rubber product and water into the second accumulating section 18C
at a rate of 1.29 wt %. Through the sight glass in the separating
vessel 4 it was observed, that the concentrated fluid H was still a
free flowing fluid. Through the sight glass at the last vent port
of the extruder 15C it could be observed that the rubber had
already changed to a crumbly state in the extruder. The crumbs
appeared white and were permanently drawn in and kneaded by the
action of the screw shafts. At the outlet section 22 rubber crumbs
of a size of roughly 2-6 mm were formed. The final bromobutyl
rubber product P collected at the outlet section was analyzed to
determine the hexane and total volatiles concentration.
Total volatiles: 0.140 wt %
Hexane: 0.055 wt %
Water: 0.085 wt %.
[0283] The process conditions and results are summarized in tables
4a), b) and c).
TABLE-US-00008 TABLE 4a Concentration Units Process conditions in
the Concentration Units P [hPa] in P [hPa] in T [.degree. C.] at
degassing T [.degree. C.] at degassing Throughput* heater vessel
heater vessel Example [kg/h] 2A** 4A 2B** 4B 20 4.4 139 756 123 130
21 4.4 157 869 147 130 22 4.6 157 796 147 130 23 4.2 157 791 147
129 *calculated on bromobutyl rubber product leaving the extruder
unit **temperature set for the heating medium
TABLE-US-00009 TABLE 4B Extruder Unit Process conditions in the
Extruder Unit* P [hPa] P [hPa] P [hPa] Stripping Stripping at vent
at vent at vent agent at Amount agent at Amount Example port 13
port 15A port 15B section 18B [wt %]** section 18C [wt %]** 20 756
270 40 water water 21 869 270 40 water 2.73 water 2.73 22 796 140
40 water 1.29 water 1.29 23 791 140 40 nitrogen 0.89 water 1.42
*The barrel temperature of the extruder was set to 30.degree. C. in
all examples **wt % with respect to rubber mass flow
TABLE-US-00010 TABLE 4c Results Contents of volatiles in the final
product Hexane Water Total Volatiles Example [wt %] [wt %]* [wt %]
20 0.08 0.12 0.20 21 0.09 0.17 0.26 22 0.10 0.08 0.18 23 0.06 0.08
0.14
[0284] The reference numerals used hereinbefore are summarized
below:
TABLE-US-00011 1 pump 2, 2A, 2B heater 3 pressure control device 4,
4A, 4B degassing vessel 4.1, 4.1A, 4.1B vapor line 4.2, 4.2A, 4.2B
pump 6 reheating unit 7 pressure control device 12 feeding point 13
rear vent port (upstream) 13.1 vapor line 15, 15A, 15B, 15B, 15C
vent port (downstream) 15.1, 15.1A, 15.1B, 15.1C vapor line 16,
16A, 16B, 16B, 16C conveying section (downstream) 18, 18A, 18B,
18B, 18C accumulating section 20 last accumulating section 22
outlet section 25 heater 26, 26A, 26B separating vessel 27, 27A,
27B aqueous phase 28, 28A, 28B organic phase 30, 30A, 30B mixing
section 32, 32A, 32B mixer 34, 34A, 34B separating wall 36, 36A,
36B recirculation pump 38, 38A, 38B recirculation line 39, 39A, 39B
coalescer 40 recirculation pump 42 recirculation line 44 heater 100
heat exchanger 200 feed cooler 300 catalyst preparation unit 400
polymerization reactor 500 mixer 600 column 700 halogenation
reactor 800 finishing equipment 900 monomer recovery unit 1000
monomer purification section 1100 solvent recovery 1200 solvent
purification section A crude fluid A C waste water D aqueous phase
for recycling E fresh water F fluid F FM feed of monomers FM.sub.R
monomer recycling stream FM.sub.R G heated fluid H H concentrated
fluid H HAL halogenating agent HBR halogenated butyl rubber BR
butyl rubber J pre-concentrated fluid J K reheated pre-concentrated
fluid K L reheated concentrated fluid L OX oxidizing agent P
product Q stopping agent S common aliphatic medium SR common
aliphatic medium recycling stream
* * * * *