Note: Descriptions are shown in the official language in which they were submitted.
ETHYLENE OLIGOMERIZATION PROCESS WITH
ALUMINOXANE/ORGANOBORON AS ACTIVATORS USING A HALOGENATED
AROMATIC SOLVENT
FIELD OF THE INVENTION
This invention relates to a novel activation system for the catalytic
oligomerization of ethylene.
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BACKGROUND OF THE INVENTION
Alpha olefins are commercially produced by the oligomerization of ethylene in
the presence of a simple alkyl aluminum catalyst (in the so called "chain
growth"
process) or alternatively, in the presence of an organometallic nickel
catalyst (in the so
called Shell Higher Olefins, or "SHOP" process). Both of these processes
typically
produce a crude oligomer product having a broad distribution of alpha olefins
with an
even number of carbon atoms (i.e. butene-1, hexene-1, octene-1, etc.). The
various
alpha olefins in the crude oligomer product are then typically separated in a
series of
distillation columns. Butene-1 is generally the least valuable of these
olefins as it is
also produced in large quantities as a by-product in various cracking and
refining
processes. Hexene-1 and octene-1 often command comparatively high prices
because
these olefins are in high demand as comonomers for linear low density
polyethylene
(LLDPE).
Technology for the selective trimerization of ethylene to hexene-1 has been
recently put into commercial use in response to the demand for hexene-1. The
patent
literature discloses catalysts which comprise a chromium source and a
pyrrolide ligand
as being useful for this process ¨ see, for example, United States Patent
("USP")
5,198,563 (Reagen et al., assigned to Phillips Petroleum).
Another family of highly active trimerization catalysts is disclosed by Wass
et al.
in WO 02/04119 (now United States Patents 7,143,633 and 6,800,702). The
catalysts
disclosed by Wass et al. are formed from a chromium source and a chelating
diphosphine ligand. These catalysts preferably comprise a diphosphine ligand
in which
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both phosphine atoms are bonded to two phenyl groups that are each substituted
with
an ortho-methoxy group. Hexene-1 is produced with high activity and high
selectivity
by these catalysts.
Similar diphosphine/tetraphenyl ligands are disclosed by Blann et at. in
W004/056478 and WO 04/056479 (now US 2006/0229480 and US 2006/0173226).
However, in comparison to the ligands of Wass et at., the
disphosphine/tetraphenyl
ligands disclosed by Blann et al. generally do not contain polar substituents
in ortho
positions. The "tetraphenyl" diphosphine ligands claimed in the '480
application must
not have ortho substituents (of any kind) on all four of the phenyl groups and
the
"tetraphenyl" diphosphine ligands claimed in '226 are characterized by having
a polar
substituent in a meta or para position.
The above described chromium/diphosphine catalysts generally require an
activator or catalyst in order to achieve meaningful rates of oligomerization.
Aluminoxane are well known activators for this catalyst system.
Methylaluminoxane
("MAO") ¨ which is made from trimethyl aluminum (TMA) - is generally preferred
in
terms of activity but suffers from a cost disadvantage.
The use of organoboron activators (as an alternative to MAO) is also known.
Such activators may also be referred to as "stoichiometric" activators because
they can
be used in essentially equimolar amounts to the chromium catalyst (i.e.
boron/chromium mole ratio of about 1/1). However, as reported in WO
2010/092554
(Kolthammer et al.), low productivity was observed where certain borate
activators were
used in an aromatic solvent ¨ a problem that was resolved by Kolthammer et al.
through the use of an aliphatic solvent. In addition, Kolthammer et at.
reported the
formation of large amounts of undesirable polymer (as a by-product) when
certain
borates were used.
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Moreover, these organoborates are typically reported to be used in combination
with an aluminum alkyl and the reactivity of the catalyst system has been
reported to be
dependent upon the amount of aluminum alkyl (McGuinness et at.,
Organomettallics,
2007, 26, 1108-1111).
Additionally, we have observed significant difficulty when attempting to
"light off'
or initiate oligomerization reactions when using organoboron activators. Many
"complete failures" were encountered ¨ i.e. the reaction produced little or no
oligomer.
We now have surprisingly discovered that these problems can be mitigated by
the process of the present invention. The process of this invention reduces
the
difficulties associated with starting up the oligomerization reaction and also
allows the
use of an organboron co-activator. The organoboron activator allows for cost
reduction
(as it allows the amount of oligomer product being produced to be increased
without
further increasing the amount of the expensive MAO co-catalyst).
SUMMARY OF THE INVENTION
In one embodiment, the present invention provides a process for the
oligomerization of ethylene, said process comprising
A) a start up step wherein ethylene is contacted with
1) an oligomerization catalyst comprising
1.1) a ligand defined by the formula (R1)(R2)-P1-bridge-P2(R3)(R4)
wherein R1, R2,R3 and R4are independently selected from the group consisting
of hydrocarbyl and heterohydrocarbyl and the bridge is a divalent moiety that
is
bonded to both phosphorus atoms;
1.2) a source of chromium that coordinates to said ligand; and
1.3) an aluminoxane,
wherein said start up step is conducted in a halogenated aromatic solvent; and
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B) a second step wherein an organoboron activator is added to said process,
wherein said second step occurs subsequent to said start up step.
Thus, for clarity: the process is started using an aluminoxane activator (but
without the organoboron activator).
The initial step of the process generally continues until "steady state" is
observed. The term "steady state" is meant to convey its conventional meaning,
namely
that the process is not proceeding in an erratic/unsteady manner. Examples of
non-
steady state conditions include rapid temperature excursions (e.g. A strong
exotherm
caused by a rapid initiator) and rapid changes in ethylene flow (in processes
where
ethylene is fed on demand, in response to changes in reactor pressure).
As a general guideline, we have observed "steady state" conditions to develop
within 5 ¨ 10 minutes of starting the reaction (under lab conditions). At this
point, the
organoboron activator is added.
DETAILED DESCRIPTION OF THE PREFERRED EMBODIMENTS
PART A: CATALYST SYSTEM
The catalyst system used in the process of the present invention must contain
four essential components, namely:
(i) a diphosphine ligand;
(ii) a source of chromium that coordinates to the ligand;
(iii) an aluminoxane; and
(iv) an organoboron activator.
Preferred forms of each of these components are discussed below.
(i) Diphosphine Liqand Used in the Oliqomerization Process
In general, the diphosphine ligand used in the oligomerization process of this
invention is defined by the formula (R1)(R2.--1_
) bridge-
P2(R3)(R4) wherein R1, R2,R3 and
R4 areindependently selected from the group consisting of hydrocarbyl and
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heterohydrocarbyl and the bridge is a divalent moiety that is bonded to both
phosphorus atoms.
The term hydrocarbyl as used herein is intended to convey its conventional
meaning ¨ i.e. a moiety that contains only carbon and hydrogen atoms. The
hydrocarbyl moiety may be a straight chain; it may be branched (and it will be
recognized by those skilled in the art that branched groups are sometimes
referred to
as "substituted"); it may be saturated or contain unsaturation and it may be
cyclic.
Preferred hydrocarbyl groups contain from 1 to 20 carbon atoms. Aromatic
groups ¨
especially phenyl groups ¨ are especially preferred. The phenyl may be
unsubstituted
(i.e. a simple C6H5 moiety) or contain substituents, particularly at an ortho
(or "o")
position.
Similarly, the term heterohydrocarbyl as used herein is intended to convey its
conventional meaning ¨ more particularly, a moiety that contains carbon,
hydrogen and
at least one heteroatom (such as 0, N, R and S). The heterohydrocarbyl groups
may
be straight chain, branched or cyclic structures. They may be saturated or
contain
unsaturation. Preferred heterohydrocarbyl groups contain a total of from 2 to
20 carbon
+ heteroatoms (for clarity, a hypothetical group that contains 2 carbon atoms
and one
nitrogen atom has a total of 3 carbon + heteroatoms).
It is preferred that each of R1, R2, R3 and R4 is a phenyl group (with an
optional
substituent in an ortho position on one or more of the phenyl groups). Highly
preferred
ligands are those in which R1 toR4 are independently selected from the group
consisting of phenyl, o-methylphenyl (i.e. ortho-methylphenyl), o-ethylphenyl,
o-isopropylphenyl and o-fluorophenyl. It is especially preferred that none of
R1 to R4
contains a polar substituent in an ortho position. The resulting ligands are
useful for the
selective tetramerization of ethylene to octene-1 with some co-product hexene
also
being produced. The term "bridge" as used herein with respect to the ligand
refers to a
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divalent moiety that is bonded to both of the phosphorus atoms in the ligand ¨
in other
words, the "bridge" forms a link between P1 and P2. Suitable groups for the
bridge
include hydrocarbyl and an inorganic moiety selected from the group consisting
of
N(CH3)-N(CH3)-, _B(R6)_, _Si(R6)2_, _p(R6)_ or _N(R6)_ where R6 is selected
from the
group consisting of hydrogen, hydrocarbyl and halogen.
It is especially preferred that the bridge is -N(R6)- wherein R5 is selected
from the
group consisting of hydrogen, alkyl, substituted alkyl, aryl, substituted
aryl, aryloxy,
substituted aryloxy, halogen, alkoxycarbonyl, carbonyloxy, alkoxy,
aminocarbonyl,
carbonylamino, dialkylamino, silyl groups or derivatives thereof and an aryl
group
substituted with any of these substituents. A highly preferred bridge is amino
isopropyl
(i.e. when R5 is isopropyl).
In one embodiment, two different types of ligands are used to alter the
relative
amounts of hexene and octene being produced. For clarity: the use of a ligand
that
produces predominantly hexene may be used in combination with a ligand that
produces predominantly octene.
(ii) Chromium Source
Any source of chromium that coordinates to the ligand and which allows the
oligomerization process of the present invention to proceed may be used.
Preferred
chromium sources include chromium trichloride; chromium (Ill) 2-
ethylhexanoate;
chromium (III) acetylacetonate and chromium carbonyl complexes such as
chromium
hexacarbonyl. It is preferred to use very high purity chromium compounds as
these
should generally be expected to minimize undesirable side reactions. For
example,
chromium acetylacetonate having a purity of higher than 99% is commercially
available
(or may be readily produced from 97% purity material ¨ using recrystallization
techniques that are well known to those skilled in the art).
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Catalyst systems comprising the above described ligands and a source of
chromium are well known for the oligomerization of ethylene. The chromium
concentrations that are typically disclosed in the relevant prior art are
generally from 20
to 400 micromolar. The present invention preferably uses a lower chromium
concentration of from 0.5 to 8 micromolar, especially from 0.5 to 5 micromolar
(i.e. from
0.5 to 8 x 10-6gram moles per litre).
(iii) Aluminoxane
Aluminoxanes are well known, commercially available items of commerce. They
may be prepared by the controlled addition of water to an alkyl aluminum
compound
such as TMA or TIBAL. Non-hydrolytic techniques to prepare aluminoxanes are
also
reported in the literature and are believed to be used by the AKZO Nobel
Company to
produce certain commercial products.
The use of methylaluminoxane (MAO) is preferred. It will be recognized by
those skilled in the art that some commercially available MAO may be made
using both
of TMA and a higher alkyl aluminum (such as TIBAL) as starting materials in
order to
improve the solubility of the resulting MAO (in comparison to a MAO made
solely from
TMA). Those MAO's are generally referred to as "modified MAO's" and they are
suitable for use in this invention.
It will also be recognized that commercially available MAO typically contains
some "residual" or "free" TMA that is associated with the MAO. "Free TMA"
typically is
present in amounts of from 10 to 40 mole ')/0 of the total aluminum contained
in the
MAO (+ TMA) and this is a preferred level for use in this invention.
Both of the TMA and MAO are expensive materials. By comparison, the current
commercial price of triethylaluminum (TEAL) is less than half of TMA or MAO
(on the
basis of cost per unit weight of aluminum). It is known that some TEAL may be
included with the MAO activator to reduce costs.
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When used, the amount of TEAL is sufficient to provide from about 10 to 70% of
the total aluminum that is added to the process on a molar basis ¨ i.e.: (the
moles of
aluminum contained in TEAL) + (the moles of aluminum contained in TEAL + TMA +
MAO) x 100% is from 10 to 70 c/o.
The amount of aluminoxane, TMA and TEAL is preferably sufficient to provide a
total Al:Cr molar ratio of from 50:1 to 1000:1, especially from 100:1 to
500:1.
It is also preferred that the aluminum concentration in the reactor is at
least 2 millimolar
(2000 micromolar) because lower levels of aluminum may not be sufficient to
"scavenge" impurities.
(iv) Organoboron Activator
Whilst not wishing to be bound by any theory, it is thought by those skilled
in the
art that the organoboron activators used in this invention provide a bulky,
labile, non-
coordinating anion which stabilizes the catalyst in a cationic form. The
bulky, non-
coordinating anion permits oligomerization to proceed at the cationic catalyst
center
(presumably because the non-coordinating anion is sufficiently labile to be
displaced by
ethylene which coordinates to the catalyst. Preferred organoboron activators
are
described in (i) ¨ (iii) below:
(i) compounds of the formula [R5] [B(R7)4]- wherein B is a boron
atom, R5 is
an aromatic hydrocarbyl (e.g. triphenyl methyl cation) and each R7 is
independently selected from the group consisting of phenyl radicals which
are unsubstituted or substituted with from 3 to 5 substituents selected
from the group consisting of a fluorine atom, a Ci-4alkyl or alkoxy radical
which is unsubstituted or substituted by a fluorine atom; and a silyl radical
of the formula -Si-(R9)3; wherein each R9 is independently selected from
the group consisting of a hydrogen atom and a Ci_4 alkyl radical; and
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(ii) compounds of the formula [(R8)t ZH][B(R7)4 wherein B is a boron atom,
H is a hydrogen atom, Z is a nitrogen atom or phosphorus atom, t is 2 or 3
and R8 is a hydrocarbyl or heterohydrocarbyl and is preferably selected
from the group consisting of C1_20 alkyl radicals, a phenyl radical which is
unsubstituted or substituted by up to three C1.4 alkyl radicals, or one R8
taken together with the nitrogen atom may form an anilinium radical and
R7 is as defined above; and
(iii) compounds of the formula B(R7)3 wherein R7 is as defined above.
In the above compounds preferably R7 is a pentafluorophenyl radical, Z is a
nitrogen atom and R8 is a C1_8 alkyl radical or R8 taken together with the
nitrogen atom
forms an anilinium radical which is substituted by two C1-4 alkyl radicals.
Examples of organobron activators include:
triethylammonium tetra(phenyl)boron,
tripropylammonium tetra(phenyl)boron,
tri(n-butyl)ammonium tetra(phenyl)boron,
trimethylammonium tetra(p-tolyl)boron,
trimethylammonium tetra(o-tolyl)boron,
tributylammonium tetra(pentafluorophenyl)boron,
tripropylammonium tetra(o,p-dimethylphenyl)boron,
tributylammonium tetra(m,m-dimethylphenyl)boron,
tributylammonium tetra(p-trifluoromethylphenyl)boron,
tributylammonium tetra(pentafluorophenyl)boron,
tri(n-butyl)ammonium tetra(o-tolyl)boron,
N,N-dimethylanilinium tetra(phenyl)boron,
N,N-diethylanilinium tetra(phenyl)boron,
N,N-diethylanilinium tetra(phenyl)n-butylboron,
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N,N-2,4,6-pentamethylanilinium tetra(phenyl)boron,
di-(isopropyl)ammonium tetra(pentafluorophenyl)boron,
dicyclohexylammonium tetra(phenyl)boron,
triphenylphosphonium tetra(phenyl)boron,
tri(methylphenyl)phosphonium tetra(phenyl)boron,
tri(dimethylphenyl)phosphonium tetra(phenyl)boron,
tropillium tetrakispentafluorophenyl borate,
triphenylmethylium tetrakispentafluorophenyl borate,
benzene (diazonium) tetrakispentafluorophenyl borate,
tropillium phenyltrispentafluorophenyl borate,
triphenylmethylium phenyltrispentafluorophenyl borate,
benzene (diazonium) phenyltrispentafluorophenyl borate,
tropillium tetrakis (2,3,5,6-tetrafluorophenyl) borate,
triphenylmethylium tetrakis (2,3,5,6-tetrafluorophenyl) borate,
benzene (diazonium) tetrakis (3,4,5-trifluorophenyl) borate,
tropillium tetrakis (3,4,5-trifluorophenyl) borate,
benzene (diazonium) tetrakis (3,4,5-trifluorophenyl) borate,
tropillium tetrakis (1,2,2-trifluoroethenyl) borate,
triphenylmethylium tetrakis (1,2,2-trifluoroethenyl) borate,
benzene (diazonium) tetrakis (1,2,2-trifluoroethenyl) borate,
tropillium tetrakis (2,3,4,5-tetrafluorophenyl) borate,
triphenylmethyliurn tetrakis (2,3,4,5-tetrafluorophenyl) borate, and
benzene (diazonium) tetrakis (2,3,4,5-tetrafluorophenyl) borate.
Readily commercially available ionic activators include:
N,N- dimethylaniliumtetrakispentafluorophenyl borate,
triphenylmethylium tetrakispentafluorophenyl borate, and
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trispentafluorophenyl borane.
PART B: PROCESS CONDITIONS
The process of this invention must be initiated in a halogenated aromatic
solvent.
Chlorinated or fluorinated solvents are preferred. It is particularly
preferred to use a
chlorobenzene solvent. Examples include monochlorobenzene, dichlorobenzene,
trichlorobenzene, and mixtures thereof (It will be appreciated by those
skilled in the art
that several isomers of dichlorobenzene and trichlorobenzene exist ¨ depending
upon
the placement of the chlorine substituent. The present invention is not
restricted to the
use of any particular such isomer).
The process of the present invention is preferably conducted in a continuously
stirred tank reactor (CSTR). The process may be batch, semi-batch, or
continuous. A
batch process requires that the reactor is not liquid full at the start of the
reaction (to
allow for the volume of product that is produced). In such a batch process,
the reactor is
preferably filled to a level of from about 10 t 25% by volume with halogenated
aromatic
solvent before the reaction is initiated.
The chromium and ligand may be present in any molar ratio which produces
oligomer, preferably between 100:1 and 1:100, and most preferably from 10:1 to
1:10,
particularly 3:1 to 1:3. Generally the amounts of (i) and (ii) are
approximately equal, i.e.
a ratio of between 1.5:1 and 1:1.5.
The catalyst components may be mixed together in the oligomerization reactor,
or ¨ alternatively ¨ some or all of the catalyst components may be mixed
together
outside of the oligomerization reactor. Suitable methods of catalyst synthesis
are
illustrated in the examples.
The aluminoxane and organoboron activators are preferably added as solutions.
In a preferred embodiment, these solutions are also prepared with a
halogenated
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aromatic solvent, although this is not essential. For clarity: the halogenated
aromatic
solvent is essential at start up, but it is not essential to use halogenated
aromatic
solvent when providing the catalyst or activator.
Ethylene is preferably fed to the reactor on demand. The ethylene may be fed
as
a gas or as solution. Again, halogenated aromatic solvent may be used to add
the
ethylene but this is not necessary.
If additional halogenated aromatic solvent is not added to the reactor during
a
continuous process, the concentration of such solvent will decrease with time
(as
solvent and product are removed from the reactor). It is generally desirable
to maintain
a concentration of halogenated solvent of at least 5 volume %.
A variety of methods are known to purify solvents used in the oligomerization
process including use of molecular sieves (3A), adsorbent alumina and
supported
de-oxo copper catalyst. Several configurations for the purifier system are
known and
depend on the nature of the impurities to be removed, the purification
efficiency
required and the compatibility of the purifier material and the process
solvent. In some
configurations, the process solvent is first contacted with molecular sieves,
followed by
adsorbent alumina, then followed by supported de-oxo copper catalyst and
finally
followed by molecular sieves. In other configurations, the solvent is first
contacted with
molecular sieves, followed by adsorbent alumina and finally followed by
molecular
sieves. In yet another configuration, the solvent is contacted with adsorbent
alumina.
Suitable temperatures range from 10 C to + 300 C preferably from 10 C to
100 C, especially from 20 to 80 C. Suitable pressures are from atmospheric to
800
atmospheres (gauge) preferably from 5 atmospheres to 100 atmospheres,
especially
from 10 to 50 atmospheres.
Irrespective of the process conditions employed, the oligomerization is
typically
carried out under conditions that substantially exclude oxygen, water, and
other
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materials that act as catalyst poisons. In addition, the reactor is preferably
purged with
a nonreactive gas (such as nitrogen or argon) prior to the introduction of
catalyst. A
purge with a solution of MAO and/or aluminum alkyl may also be employed to
lower the
initial level of catalyst poisons. The use of hydrogen is especially preferred
because it
has been observed to reduce the amount of polymer that is formed. The most
preferred catalysts of this invention predominantly produce octene with some
hexene
(as shown in the examples) but smaller quantities of butene and Ci0+ olefins
are also
produced. The crude product stream may be separated into various fractions
using, for
example, a conventional distillation system. It is within the scope of this
invention to
recycle the "whole" oligomer product or some fraction(s) thereof to the
reaction for use
as an oligomerization diluent. For example, by recycling a butene rich stream
it might
be possible to lower the refrigeration load in distillation. Alternatively,
the C10+ fraction
might be preferentially recycled to improve the solubility of one or more
components of
the catalyst system.
Techniques for varying the distribution of products from the oligomerization
reactions include controlling process conditions (e.g. concentration of
components (i)-
(iii), reaction temperature, pressure, residence time) and properly selecting
the design
of the process and are well known to those skilled in the art.
In another embodiment, a catalyst that produces ethylene homopolymer is
deliberately added to the reactor in an amount sufficient to convert from 1 to
5 weight%
of the ethylene feed to an ethylene homopolymer. This catalyst is preferably
supported.
The purpose is to facilitate the removal of by-product polyethylene.
The ethylene feedstock for the oligomerization may be substantially pure or
may
contain other olefinic impurities and/or ethane. One embodiment of the process
of the
invention comprises the oligomerization of ethylene-containing waste streams
from
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other chemical processes or a crude ethylene/ethane mixture from an ethane to
ethylene cracker.
The feedstock is preferably treated to remove catalyst poisons (such as
oxygen,
water and polar species) using techniques that are well known to those skilled
in the
art. The technology used to treat feedstocks for polymerizations is suitable
for use in
the present invention and includes the molecular sieves, alumina and de-oxo
catalysts
described above for analogous treatment of the process solvent.
Reactor Systems
A general review of suitable reactors for selective oligomerization is
provided
first, followed by a detailed description of preferred reactor designs. There
exist a
number of options for the oligomerization reactor including batch, semi-batch,
and
continuous operation. Oligomerization reactions can generally be performed
under a
range of process conditions that are readily apparent to those skilled in the
art.
Evaporative cooling from one or more monomers or inert volatile liquids is but
one (prior
art) method that can be employed to effect the removal of heat from the
reaction. The
reactions may be performed in the known types of reactors, such as a plug-flow
reactor,
or a continuously stirred tank reactor (CSTR), or a loop reactor, or
combinations
thereof. A wide range of methods for effecting product, reactant, and catalyst
separation and/or purification are known to those skilled in the art and may
be
employed: distillation, filtration, liquid-liquid separation, slurry settling,
extraction, etc.
One or more of these methods may be performed separately from the
oligomerization
reaction or it may be advantageous to integrate at least some with the
reaction; a non-
limiting example of this would be a process employing catalytic (or reactive)
distillation.
Also advantageous may be a process which includes more than one reactor, a
catalyst
kill system between reactors or after the final reactor, or an integrated
reactor/separator/purifier. While all catalyst components, reactants, inerts,
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products could be employed in the present invention on a once-through basis,
it is often
economically advantageous to recycle one or more of these materials; in the
case of
the catalyst system, this might require reconstituting one or more of the
catalysts
components to achieve the active catalyst system.
More specific reactor designs have been described in the patent literature:
= a liquid phase reactor with "bubbling" ethylene feed is taught as a means
to mitigate PE formation (WO 2009/060342, Kleingeld et al.);
= a liquid phase reactor with an inert, condensable liquid is claimed as a
means to improve temperature control (WO 2009/060343, Crildenhuys). The
condensable liquid boils from the reaction liquid and is condensed overhead;
and
= the use of a liquid/gas phase reactor in which cooling coils are present
in
the gas phase head space is described in WO 2007/016996, Fritz et al.).
The present invention also encompasses other reactor designs for selective
oligomerizations. A non adiabatic reactor system may be used. The term "non
adiabatic" means that heat is added to and/or removed from the oligomerization
reactor. The term "reactor system" means that one or more reactors are
employed
(and the term "non adiabatic reactor system" means that at least one of the
reactors is
equipped with a heat exchanger that allows heat to be added to or removed from
it).
One design relates to a CSTR with an external heat exchanger. A second design
relates to a tubular plug flow equipped with multiple feed ports for ethylene
along the
length of the reactor. A third design relates to a combination of a CSTR
followed by a
tubular reactor. A fourth design provides a loop reactor. A fifth design
provides a
reactor having an internal cooling system (such as a draft tube reactor).
One preferred CSTR for use in the present invention is equipped at least one
external heat exchanger ¨ meaning that the heat exchanger surface(s) are not
included
within the walls of the CSTR. The term "heat exchanger" is meant to include
its broad,
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conventional meaning. Most importantly, the heat exchanger will preferably be
designed so as to allow heating of the reactor contents (which may be
desirable during
start up) and to provide heat removal during the oligomerization. A preferred
external
heat exchanger for a CSTR comprises a conventional shell and tube exchanger
with a
"process" side tube system and a shell for the exchange side. In one
embodiment the
"process side" (i.e. the side of the exchanger that contains the fluid from
the
oligomerization process) is a tube that exits the reactor and flows through
the shell for
heat exchange, then reenters the reactor with cooled (or heated) process
fluid. For
clarity: during an oligomerization reaction a portion of the hot reactor
contents or
"process fluid" will flow from the reactor to the external heat exchanger,
through a tube.
The exterior of the tube comes into contact with cold fluid on the shell side
of the
exchanger, thus cooling the process fluid. The cooled process fluid is then
returned to
the reactor.
The use of two of more CSTR reactors in series is also contemplated. In
particular, the use of a first CSTR having a small volume followed by a larger
CSTR
might be used to facilitate startup.
In another embodiment, a heat exchanger is located between two CSTRs. In
this embodiment, the product from the first oligomerization reactor leaves
that reactor
through an exit tube. The oligomerization products in this exit tube are then
directed
through a heat exchanger. After being cooled by the heat exchanger, the
oligomerization products are then directed into a second CSTR. Additional
ethylene
(and, optionally, catalyst) is added to the second CSTR and further
oligomerization
takes place. '
The amount of heat generated by the oligomerization reaction is generally
proportional to the amount of ethylene being oligomerized. Thus, at high rates
of
oligomerization, a high rate of coolant flow is required in the shell side of
the exchanger.
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The rate of oligomerization is generally proportional to the amount of
ethylene
and catalyst that are fed to the CSTR. In one preferred embodiment the
ethylene is first
contacted with solvent in a mixing vessel that is external to the CSTR. For
convenience, this mixing vessel is referred to herein as a "solution
absorber". The
solution absorber is preferably equipped with a heat exchanger to remove the
heat of
absorbtion ¨ i.e. heat is generated when the ethylene dissolves in the solvent
and this
heat exchanger removes the heat of solution. The solution absorber may be a
CSTR,
or alternatively, a simple plug flow tube. Thus, the heat exchanger on the
solution
absorber is used to provide cooled feed. In one embodiment the heat exchanger
may
be used to chill the feed to below ambient conditions ¨ this is desirable to
maximize
reactor throughput.
In a preferred design, another heat exchanger is provided that allows the feed
stream to be heated. This heat exchanger may be located in direct contact with
the
solution absorber or ¨ alternatively, this heat exchanger may be located
between the
solution absorber and the oligomerization reactor. In general, this heat
exchanger will
be used during non-steady state conditions (such as are encountered at start
up or
during a reactor upset) to quickly provide heat to the reactor.
The CSTR is preferably operated in continuous flow mode ¨ i.e. feed is
continuously provided to the CSTR and product is continuously withdrawn.
The CSTR described above may be used to provide the high degree of
temperature control that we have observed to be associated with a low degree
of
polymer formation.
In another design, the CSTR is equipped with one or more of the mixing
elements described in USP 6,319,996 (Burke et al.). In particular, Burke et
al. disclose
the use of a tube which has a diameter that is approximately equal to the
diameter of
the agitator of the CSTR. This tube extends along the length of the agitator
shaft,
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thereby forming a mixing element that is often referred to as a "draft tube"
by those
skilled in the art. The reactor used in this invention may also employ the
mixing helix
disclosed by Burke et al. (which helix is located within the draft tube and
forms a type of
auger or Archimedes screw within the draft tube). The use of stationary,
internal
elements (to divide the CSTR into one or more zones) may also be employed. In
one
such example, two impellers are vertically displaced along the length of the
agitation
shaft i.e. one in the top part of the reactor and another in the bottom. An
internal "ring"
or "doughnut" is used to divide the CSTR into a top reaction zone and a bottom
reaction
zone. The ring is attached to the diameter of the CSTR and extends inwardly
towards
the agitation shaft to provide a barrier between the top and bottom reaction
zones. A
hole in the center of the ring allows the agitation shaft to rotate freely and
provides a
pathway for fluid flow between the two reactions zones. The use of such rings
or
doughnuts to divide a CSTR into different zones is well known to those skilled
in the art
of reactor design.
In another design, two or more separate agitators with separate shafts and
separate drives may be employed. For example, a small impeller might be
operated at
high velocity/high shear rate to disperse the catalyst and/or ethylene as it
enters the
reactor and a separate (larger) impeller with a draft tube could be used to
provide
circulation within the reactor.
An alternative reactor design is a tubular/plug flow reactor with an external
heat
exchanger. Tubular/plug flow reactors are well known to those skilled in the
art. In
general, such reactors comprise one or more tubes with a length/diameter ratio
of from
10/1 to 1000/1. Such reactors are not equipped with active/powered agitators
but may
include a static mixer. Examples of static mixers include those manufactured
and sold
by Koch-Glitsch Inc. and Sulzer-Chemtech.
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Tubular reactors for use in the present invention are preferably characterized
by
two features:
1) external cooling; and
2) the use of at least one incremental ethylene feed port along the length
of
the tubular reactor (i.e. in addition to the initial ethylene feed at the
start of the tubular
reactor).
In one embodiment, the tubular reactor is a so called "heat-exchange reactor"
which is generally configured as a tube and shell heat exchanger. The
oligomerization
reaction occurs inside the tube(s) of this reactor. The shell side provides a
heat
exchange fluid (for the purposes described above, namely to heat the reaction
during
start up and/or to cool the reaction during steady state operations).
In one embodiment, the tubes are bent so as to form a type of static mixer for
the
fluid passing through the shell side. This type of heat exchanger is known to
those
skilled in the art and is available (for example) from Sulzer-Chemtech under
the trade
name SMR.
It is especially preferred that the Reynolds number of the reaction fluid that
flows
through the tube (or tubes) of the tubular reactor is from 2,000 to
10,000,000. Reynolds
number is a dimensionless number that is readily calculated using the
following
formula:
Re = pVL
where:
V is the mean fluid velocity (SI units: m/s);
L is a characteristic linear dimension (e.g. internal diameter of tube);
p is the dynamic viscosity of the fluid (Pa.s or II=s/m2 or kg/(m.$)); and
p is the density of the fluid (kg/m3).
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In one such embodiment a plurality of heat exchange reactors are connected in
series. Thus, the process flow that exits the first reactor enters the second
reactor.
Additional ethylene is added to the process flow from the first reactor but
additional
catalyst is preferably not added.
In another embodiment, a CSTR is connected in series to a tubular reactor. One
sub embodiment of this dual reactor system comprises a CSTR operated in
adiabatic
mode, followed by a tubular reactor having an external heat exchanger ¨ in
this
embodiment the amount of ethylene that is consumed (i.e. converted to
oligomer) in the
CSTR is less than 50 weight % of the total ethylene that is consumed in the
reactors.
In another sub embodiment of this dual reactor system, a CSTR that is equipped
with
an external heat exchanger is connected to a downstream tubular reactor that
is
operated in adiabatic mode. In this embodiment, the amount of ethylene that is
converted/consumed in the CSTR is in excess of 80 weight % of the ethylene
that is
consumed in the reactor. The tubular reactor may also have several different
ports
which allow the addition of catalyst killer/deactivator along the length of
the reactor. In
this manner, some flexibility is provided to allow the reaction to be
terminated before
the product exits from the reactor.
Another reactor design for use in the present invention is a loop reactor.
Loop
reactors are well known and are widely described in the literature. One such
design is
disclosed in USP 4,121,029 (Irvin et al.). The loop reactor disclosed by Irvin
et al.
contains a "wash column" that is connected to the upper leg of the loop
reactor and is
used for the collection of polymer. A similar "wash column" is contemplated
for use in
the present invention to collect by-product polymer (and/or supported
catalyst). A
hydrocyclone at the top end of the wash column may be used to facilitate
polymer
separation.
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A fifth reactor design for use in the present invention is another type of
heat
exchange reactor in which the process side (i.e. where the oligomerization
occurs) is
the "shell side" of the exchanger. One embodiment of this reactor design is a
so called
"draft tube" reactor of the type reported to be suitable for the
polymerization of butyl
rubber. This type of reactor is characterized by having an impeller located
near the
bottom of the reactor, with little or no agitator shaft extending into the
reactor. The
impeller is encircled with a type of "draft tube" that extends upwards through
the center
of the reactor. The draft tube is open at the bottom (to allow the reactor
contents to be
drained into the tube, for upward flow) and at the top ¨ where the reactor
contents are
discharged from the tube. A heat exchanger tube bundle is contained within the
reactor
and is arranged such that the tubes run parallel to the draft tube and are
generally
arranged in a concentric pattern around the draft tube. Coolant flows through
the tubes
to remove the heat of the reaction.
Monomer is preferably added by one or more feed ports that are located on the
perimeter of the reactor (especially near the bottom of the reactor) and
oligomerization
product is withdrawn through at least one product exit port (preferably
located near the
top of the reactor). Catalyst is preferably added through a separate feed line
that is not
located close to any of the monomer feed ports(s) or product exit port(s).
Draft tube
reactors are well known and are described in more detail in USP 4,007,016
(Weber)
and USP 2,474,592 (Palmer) and the references therein. Figure 2 of USP
2,474,592
illustrates the use of a fluid flushing system to flush the agitator shaft in
the vicinity of
the agitator shaft seal. More specifically, a fluid chamber through the
agitator shaft seal
is connected to a source of flushing fluid (located outside of the reactor)
and the
channel terminates in the area where the agitator shaft enters the reactor.
"Flushing
fluid" is pumped through the channel to flush the base of the agitator and
thereby
reduce the amount of polymer build up at this location.
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Another form of this type of reactor (i.e. in which the process is undertaken
on
the "shell" side of an internally heat exchanged reactor) is sold by ABB
Lummus under
the trademark Helixchanger .
Another known technique to reduce the level of fouling in a chemical reactor
is to
coat the reactor walls and/or internals and/or agitators with a low fouling
material such
as glass or polytetraflouroethylene (PTFE). The use of coatings can be
especially
beneficial on high fouling areas such as agitator shafts and impellers.
Reactor Control
The control systems required for the operation of CSTR's and tubular reactors
are well known to those skilled in the art and do not represent a novel
feature of the
present invention. In general, temperature, pressure and flow rate readings
will provide
the basis for most conventional control operations. The increase in process
temperature (together with reactor flow rates and the known enthalpy of
reaction) may
be used to monitor ethylene conversion rates. The amount of catalyst may be
increased to increase the ethylene conversion (or decreased to decrease
ethylene
conversion) within desired ranges. Thus, basic process control may be derived
from
simple measurements of temperature, pressure and flow rates using conventional
thermocouples, pressure meters and flow meters. Advanced process control (for
example, for the purpose of monitoring product selectivity or for the purpose
of
monitoring process fouling factors) may be undertaken by monitoring additional
process
parameters with more advanced instrumentation. Known/existing instrumentation
that
may be employed include in-line/on-line instruments such as NIR infrared,
Fourier
Transform Infrared (FTIR), Raman, mid-infrared, ultra violet (UV)
spectrometry, gas
chromatography (GC) analyzer, refractive index, on-line densitometer or
viscometer.
The use of NIR or GC to measure the composition of the oligomerization reactor
and
final product composition is especially preferred.
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The measurement may be used to monitor and control the reaction to achieve
the targeted stream properties including but not limited to concentration,
viscosity,
temperature, pressure, flows, flow ratios, density, chemical composition,
phase and
phase transition, degree of reaction, polymer content, selectivity.
The control method may include the use of the measurement to calculate a new
control set point. The control of the process will include the use of any
process control
algorithms, which include, but are not limited to the use of PID, neural
networks,
feedback loop control, forward loop control and adaptive control.
Catalyst Deactivation, Catalyst Removal and Polymer Removal
In general, the oligomerization catalyst is preferably deactivated immediately
downstream of the reactor as the product exits the reaction vessel. This is to
prevent
polymer formation and potential build up downstream of the reactor and to
prevent
isomerisation of the 1-olefin product to the undesired internal olefins. It is
generally
preferred to flash and recover unreacted ethylene before deactivation.
However, the
option of deactivating the reactor contents prior to flashing and recovering
ethylene is
also acceptable. The flashing of ethylene is endothermic and may be used as a
cooling
source. In one embodiment, the cooling provided by ethylene flashing is used
to chill a
feedstream to the reactor.
In general, many polar compounds (such as water, alcohols and carboxylic
acids) will deactivate the catalyst. The use of alcohols and/or carboxylic
acids is
preferred ¨ and combinations of both are contemplated. It is generally found
that the
quantity employed to deactivate the catalyst is sufficient to provide
deactivator to metal
(from activator) mole ratio between about 0.1 to about 4. The deactivator may
be
added to the oligomerization product stream before or after the volatile
unreacted
reagents/diluents and product components are separated. In the event of a
runaway
reaction (e.g. rapid temperature rise) the deactivator can be immediately fed
to the
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oligomerization reactor to terminate the reaction. The deactivation system may
also
include a basic compound (such as sodium hydroxide) to minimize isomerization
of the
products (as activator conditions may facilitate the isomerization of
desirable alpha
olefins to undesired internal olefins).
Polymer removal (and, optionally, catalyst removal) preferably follows
catalyst
deactivation. Two "types" of polymer may exist, namely polymer that is
dissolved in the
process solvent and non-dissolved polymer that is present as a solid or
"slurry".
Solid/non-dissolved polymer may be separated using one or more of the
following types of equipment: centrifuge; cyclone (or hydrocyclone), a
decanter
equipped with a skimmer or a filter. Preferred equipment include so called
"self
cleaning filters" sold under the name V-auto strainers, self cleaning screens
such as
those sold by Johnson Screens Inc. of New Brighton, Minnesota and centrifuges
such
as those sold by Alfa Laval Inc. of Richmond, VA (including those sold under
the trade
name Sharpies).
Soluble polymer may be separated from the final product by two distinct
operations. Firstly, low molecular weight polymer that remains soluble in the
heaviest
product fraction (C20+) may be left in that fraction. This fraction will be
recovered as
"bottoms" from the distillation operations (described below). This solution
may be used
as a fuel for a power generation system.
An alternative polymer separation comprises polymer precipitation caused by
the
removal of the solvent from the solution, followed by recovery of the
precipitated
polymer using a conventional extruder. The technology required for such
separation/recovery is well known to those skilled in the art of solution
polymerization
and is widely disclosed in the literature.
In another embodiment, the residual catalyst is treated with an additive that
causes some or all of the catalyst to precipitate. The precipitated catalyst
is preferably
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removed from the product at the same time as by-product polymer is removed
(and
using the same equipment). Many of the catalyst deactivators listed above will
also
cause catalyst precipitation. In a preferred embodiment, a solid sorbent (such
as clay,
silica or alumina) is added to the deactivation operation to facilitate
removal of the
deactivated catalyst by filtration or centrifugation.
Reactor fouling (caused by deposition of polymer and/or catalyst residue) can,
if
severe enough, cause the process to be shut down for cleaning. The deposits
may be
removed by known means, especially the use of high pressure water jets or the
use of
a hot solvent flush. The use of a halogenated aromatic solvent for solvent
flushing is
generally preferred because they are good solvents for polyethylene. The use
of the
heat exchanger that provides heat to the present process may also be used
during
cleaning operations to heat the cleaning solvent.
Distillation
In one embodiment of the present invention, the oligomerization product
produced from this invention is added to a product stream from another alpha
olefins
manufacturing process for separation into different alpha olefins. As
previously
discussed, "conventional alpha olefin plants" (wherein the term includes i)
those
processes which produce alpha olefins by a chain growth process using an
aluminum
alkyl catalyst, ii) the aforementioned "SHOP" process and iii) the production
of olefins
from synthesis gas using the so called Lurgi process) have a series of
distillation
columns to separate the "crude alpha product" (i.e. a mixture of alpha
olefins) into alpha
olefins (such as butene-1, hexene-1 and octene-1). The mixed hexene-octene
product
which is preferably produced in accordance with the present invention is
highly suitable
for addition/mixing with a crude alpha olefin product from an existing alpha
olefin plant
(or a "cut" or fraction of the product from such a plant) because the mixed
hexene-
octene product produced in accordance with the present invention can have very
low
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levels of internal olefins. Thus, the hexene-octene product of the present
invention can
be readily separated in the existing distillation columns of alpha olefin
plants (without
causing the large burden on the operation of these distillation columns which
would
otherwise exist if the present hexene-octene product stream contained large
quantities
of internal olefins). As used herein, the term "liquid product" is meant to
refer to the
oligomers produced by the process of the present invention which have from 4
to
(about) 20 carbon atoms.
In another embodiment, the distillation operation for the oligomerization
product
is integrated with the distillation system of a solution polymerization plant.
The liquid product from the oligomerization process of the present invention
preferably consists of from 20 to 80 weight % octenes (especially from 35 to
75
weight %) octenes and from 15 to 50 weight A) (especially from 20 to 40
weight %)
hexenes (where all of the weight % are calculated on the basis of the liquid
product by
100%.
The preferred oligomerization process of this invention is also characterized
by
producing very low levels of internal olefins (i.e. low levels of hexene-2,
hexene-3,
octene-2, octene-3 etc.), with preferred levels of less than 10 weight A
(especially less
than 5 weight %) of the hexenes and octenes being internal olefins.
EXAMPLES
The following abbreviations are used in the examples:
A = Angstrom units
NMR = nuclear magnetic resonance
Et = ethyl
Bu = butyl
iPr = isopropyl
c* = comparative
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rpm = revolutions per minute
GC = gas chromatography
R, = reaction
Wt = weight
Ca's = butenes
C8's = hexenes
C8's = octenes
PE = polyethylene
Part I: Preferred Ligand Synthesis
General
This section illustrates the synthesis of a preferred but non-limiting ligand
for use
in the present invention.
All reactions involving air and or moisture sensitive compounds were conducted
under nitrogen using standard Schlenk or cannula techniques, or in a glovebox.
Reaction solvents were purified prior to use (e.g. by distillation) and stored
over
activated 4 A sieves. Diethylamine, triethylamine and isopropylamine were
purchased
from Aldrich and dried over 4 A molecular sieves prior to use. 1-Bromo-2-
fluoro-
benzene, phosphorus trichloride (PCI3), hydrogen chloride gas and n-
butyllithium were
purchased from Aldrich and used as is. The methylalumoxane (MAO), 10 weight %
Al
in toluene, was purchased from Akzo and used as is. Deuterated solvents were
purchased (toluene-d8, THF-d8) and were stored over 4 A sieves. NMR spectra
were
recorded on a Bruker 300 MHz spectrometer (300.1 MHz for 1H, 121.5 MHz for
31P,
282.4 for 19F).
Preparation of Et2NPCI2
Et2NH (50.00 mmol, 5.17 mL) was added dropwise to a solution of PCI3 (25.00
mmol, 2.18 mL) in diethyl ether (hereinafter "ether") (200 mL) at ¨78 C. After
the
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addition, the cold bath was removed and the slurry was allowed to warm to room
temperature over 2 hours. The slurry was filtered and the filtrate was pumped
to
dryness. The residue was distilled (500 microns, 55 C) to give the product in
quantitative yield. 1H NMR (8, toluene-d8): 2.66 (doublet of a quartets, 4H,
Jph = 13
Hz, Jhh = 7 Hz), 0.75 (triplet, 6H, J = 7 Hz).
Pre aration of P-NEt
To solution of n-BuLi (17.00 mL of 1.6 M n-BuLi hexane solution, 27.18 mmol)
in
ether (100 mL) maintained at ¨85 C, was added dropwise a solution of 1-bromo-2-
fluorobenzene (4.76 g, 27.18 mmol) in ether (40 mL) over 2 hours. After
addition, the
reaction flask was stirred for 1 hour at -78 C, resulting in a white slurry.
Et2NPCI2
(2.36 g, 13.58 mmol) in ether (20 mL) was then added very slowly while the
reaction
temperature was maintained at -85 C. The reaction was allowed to warm to -10 C
overnight. Toluene (10 mL) was then added to the reaction flask and the
volatiles were
removed in vacuo. The residue was extracted with toluene and the solution was
pumped to dryness. The crude product was distilled (300 microns, 100 C)
yielding
3.78 g (95%) of product. 1H NMR (8, THF-d8): 7.40-7.01 (4 equal intense
multiplets,
8H), 3.11 (doublets of quartet, 4H, Jph = 13 Hz, Jhh = 7 Hz), 0.97 (triplet,
6H, J = 7 Hz).
19F NMR (8, THF-d8): -163.21 (doublet of multiplets, J = 48 Hz). GC-MS. M+ =
293.
Preparation of (ortho-F-C61-14)2PCI
Anhydrous HCI(g) was introduced to the head space of an ethereal solution (100
mL) of (ortho-F-C81-14)P-NEt2 (3.73 g, 12.70 mmol) to a pressure of 3 psi. A
white
precipitate formed immediately. The reaction was stirred for an additional 0.5
hours at
which point the slurry was pumped to dryness to remove volatiles. The residue
was re-
slurried in ether (100 mL) and filtered. The filtrate was pumped to dryness
yielding
(ortho-F-C81-14)2PCI as a colorless oil in quantitative yield. 1H NMR (5, THF-
d8): 7.60
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(m, 4H), 7.20 (m, 2H), 7.08 (m, 2H). 19F NMR (8, THF-d8): -106.94 (doublet of
multiplets, J = 67 Hz).
Preparation of (ortho-F-CEH4)2PNH(i-Pr)
To a solution of (ortho-F-C81-14)PCI (1.00 g, 3.90 mmol) in ether (50 mL) and
NEt3
(3 mL) was added an ethereal solution of i-PrNH2 (0.42 mL, 4.90 mmol) at -5 C.
Immediate precipitate was observed. The slurry was stirred for 3 hours and
filtered.
The filtrate was pumped to dryness to give a colorless oil of (ortho-F-C61-
14)PNH(i-Pr) in
quantitative yield. 1H NMR (8, THF-d8): 7.42 (m, 2H), 7.30 (m, 2H), 7.11 (m,
2H), 6.96
(m, 2H), 3.30 (septet, 1H, J= 7 Hz), 2.86 (br s, 1H), 1.15 (d, 6H, J = 7 Hz).
19F NMR (8,
THF-d8): -109.85 (doublet of multiplets, J = 40 Hz). GC-MS, M+ = 279.
Preparation of (ortho-F-C8F14)2PN(i-Pr)P(ortho-F-C6H4)2 ("Liqand 1")
To a solution of (ortho-F-C81-14)2PNH(i-Pr) (3.90 mmol) [made from i-PrNH2 and
(ortho-F-C81-14)2PCI (1.00 g, 3.90 mmol)] in ether (100 mL) maintained at -70
C was
added dropwise a solution of n-BuLi (2.43 mL of 1.6 M n-BuLi hexane solution,
3.90 mmol)). The mixture was stirred at -70 C for 1 hour and allowed to warm
to -10 C
in a cold bath (2 hours). The solution was re-cooled to -70 C and (ortho-F-C81-
14)2PCI
(1.00 g, 3.90 mmol) was slowly added. The solution was stirred for 1 hour at -
70 C and
allowed to slowly warm to room temperature forming a white precipitate. The
slurry
was pumped to dryness and the residue was extracted with toluene and filtered.
The
filtrate was pumped to dryness and recrystallized from heptane at -70 C (2x)
yielding
1.13 g (58%) of product. At room temperature this material was an oil which
contained
both the desired ligand (ortho-F-C61-14)2PN(i-Pr)P(ortho-F-C8H4)2 and its
isomer (ortho-
F-C8H4)2P[=NO-PIP(ortho-F-C8H4)2. A toluene solution of this mixture and 50 mg
of
(ortho-F-C81-14)2PCI was heated at 65 C for three hours to convert the isomer
to the
desired ligand. 1H NMR (THF-d8, 8): 7.35 (m, 8H), 7.10 (m, 4H), 6.96 (m, 4H),
3.94
(m, 1H), 1.24 (d, 6H, J = 7Hz). 19F NMR (THF-d8, 8): -104.2 (br. s).
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In a more preferred procedure the initial steps of the synthesis are conducted
in
pentane at -5 C (instead of ether) with 10% more of the (ortho-F-C6H4)2PCI
(otherwise
as described above). This preferred procedure allows (ortho-F-C6F14)2PN(i-
Pr)P(ortho-
F-C6H4)2 to be formed in high (essentially quantitative) yield without the
final step of
heating in toluene.
Catalyst Preparation
Part II: Oligomerization Reaction
General
The aluminoxane used in all experiments was purchased from Albemarle
Corporation and reported to contain 10 weight % aluminum. The product was
described as a conventional methylaluminoxane that was prepared using TMA as
the
only source of an aluminum (i.e., it was not a so-called "modified MAO"). The
"free
TMA" content was reported to be about 10 mole % - i.e. for every 100 moles of
aluminum in the product, 90 moles were contained in the aluminoxane oligomer
and 10
were present as "free TMA". For convenience, this product is referred to as
"MAO" in
the accompanying table and detailed experimental description. (For further
certainty:
the "Al(MA0)" column includes the aluminum contained in both the aluminoxane
oligomer and free TMA. For example, the value of 1,000 micromoles - for
inventive run
17- represents 900 micromoles of aluminum in the oligomer and 100 micromoles
of
free TMA.)
EXAMPLES
Comparative Run 1 - Baseline Run in 1-Octene; MAO only
A 600 mL reactor fitted with a stirrer was purged 3 times with argon while
heated
at 80 C. The reactor was then cooled to 55 C (-5 C below reaction temperature)
and a
solution of MAO (1.44 g, 10 weight % MAO) in 65 g of 1-octene (containing 5.97
weight
% cyclohexane as internal reference) was transferred via a stainless steel
cannula to
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CA 02758126 2011-11-08
the reactor, followed by 78 g of 1-octene (containing 5.97 weight A)
cyclohexane).
Stirrer was started and set to 1700 rpm. The reactor was then pressurized to
39 bar
with ethylene and temperature adjusted to 47 C. Ligand 1 (4.22 mg, 0.0084
mmol) and
chromium acetylacetonate (2.88 mg, 0.0082 mmol) were premixed in 14.3 g of 1-
octene
(containing 5.97 weight % cyclohexane) in a hypovial. The mixture was
transferred
under ethylene to the pressurized reactor and then the reactor pressure was
immediately increased to 45 bar with ethylene. The reaction was allowed to
proceed
for 20 minutes while maintaining the temperature at 60 C. The reaction was
terminated
by stopping ethylene flow to the reactor and cooling the contents to 30 C.
Stirring was
stopped and reactor slowly depressurized to atmospheric pressure. Reactor was
then
opened and product mixture transferred to a pre-weighed flask containing 1.5 g
of
isopropanol. The mass of product produced was 85.6 g. A sample of the liquid
product
was analyzed by GC-FID.
(Comparative) Example 2 - Baseline Run in Cvclohexane; MAO only
A 600 mL reactor fitted with a stirrer was purged 3 times with argon while
heated
at 80 C. The reactor was then cooled to 45 C (-2 C below reaction temperature)
and a
solution of MAO (1.44 g, 10 weight % MAO) in 65 g of cyclohexane was
transferred via
a stainless steel cannula to the reactor, followed by 78 g of cyclohexane.
Stirrer was
started and set to 1700 rpm. The reactor was then pressurized to 35 bar with
ethylene
and temperature adjusted to 47 C. Ligand 1 (4.43 mg, 0.0089 mmol) and chromium
acetylacetonate (3.02 mg, 0.0087 mmol) were premixed in 14.3 g of cyclohexane
in a
hypovial. The mixture was transferred under ethylene to the pressurized
reactor and
then the reactor pressure was immediately increased to 40 bar with ethylene.
The
reaction was allowed to proceed for 15 minutes while maintaining the
temperature at
46 C. The reaction was terminated by stopping ethylene flow to the reactor and
cooling
the contents to 30 C. Stirring was stopped and reactor slowly depressurized to
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CA 02758126 2011-11-08
atmospheric pressure. Reactor was then opened and product mixture transferred
to a
pre-weighed flask containing 1.5 g of isopropanol. The mass of product
produced was
100.3 g. A sample of the liquid product was analyzed by GC-FID.
Comparative Example 3 - Borate Attempts
A 600 mL reactor fitted with a stirrer was purged 3 times with argon while
heated at
80 C. The pressure was reduced to 1 bar and a solution of triethylaluminium
(0.338
mmol) in cyclohexane was added and the reactor placed under 10 bar ethylene
pressure. In a hypo vial, triethylaluminium (49.92 x10-3 mmol) in cyclohexane
was
added to a solution of chromium (0.93 x10-3 mmol) and isoheptyl-N(P(C6H5)2)2
(1.11
x10-3 mmol) in cyclohexane, and stirred for -30 seconds. Addition of
Roctadecy1)2MeNKB(C6F5)4] (1.11 x10-3 mmol) in cyclohexane was then added and
the mixture stirred for a further 1 minute. The reactor pressure was reduced
to 1 bar at
which point the activation solution was added and the reactor immediately
pressurized
to 45 bar with ethylene. The pressure was kept constant at 45 bar for between
20 - 60
minutes. The reaction was terminated by stopping ethylene flow to the reactor
and
cooling the contents to 30 C. Stirring was stopped and reactor slowly
depressurized to
atmospheric pressure. Reactor was then opened and product mixture transferred
to a
pre-weighed flask containing 1.5 g of isopropanol. There was no observable
ethylene
consumption and no measureable amount of product was produced. Several other
attempts were made in which the catalyst concentration, time for catalyst
preparation
and the sequence of addition of the catalyst components were varied; however,
the
extent of reaction was either very low or not observable.
Comparative Example 4
A 600 mL reactor fitted with a stirrer was purged 3 times with argon while
heated at
80 C. The reactor was then cooled to 45 C (-2 C below reaction temperature)
and a
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CA 02758126 2011-11-08
solution of MAO (0.44 g, 7.1 weight % Al in MMA0-3A solution in isopentane)
topped
up to 64.8 g with cyclohexane was transferred via a stainless steel cannula to
the
reactor, followed by 63.9 g of cyclohexane. Stirrer was started and set to
1700 rpm.
The reactor was then pressurized to 30 bar with ethylene, 10.30 g of a
cyclohexane
solution of isoheptyl-N(P(C6H5)2)2 (1.295 mg, 2.68 x10-3 mmol) and chromium
acetylacetonate (0.916 mg, 2.68 x10-3 mmol) was added to the reactor with an
additional 5 bar ethylene, and the temperature adjusted to 47 C. The reaction
was
allowed to proceed for 20 minutes while maintaining the temperature at 48 C
and the
pressure at 35 bar, at which point Roctadecy1)2MeNHHB(C6F5)4] (3.820 mg, 3.10
x10-3
mmol) in cyclohexane was added to the reactor. The reaction was allowed to
proceed
for an additional 15 minutes and then terminated by stopping ethylene flow to
the
reactor and cooling the contents to 30 C. Stirring was stopped and reactor
slowly
depressurized to atmospheric pressure. Reactor was then opened and product
mixture
transferred to a pre-weighed flask containing 1.5 g of isopropanol. The mass
of product
produced was 99.2g. A sample of the liquid product was analyzed by GC-FID.
The addition of the borate activator produced an exotherm, indicating a sudden
increase in reaction rate. This was confirmed by the ethylene flow rate to the
reactor,
which showed an increase of more than 50%.
Applicants believe this result to be novel and surprising. However, this
example
is comparative as we have discovered that further improvements may be obtained
through the use of a halogenated aromatic solvent as shown below.
The above described example was conducted as a batch reaction. The reactor
quickly became full of product so the reaction was terminated.
A continuous reaction or a semi-batch reaction may also be conducted using the
above described start up protocol (i.e. initializing the reaction with MAO
only). As will be
appreciated by those skilled in the art, a continuous reaction is one in which
reactants
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CA 02758126 2011-11-08
are continuously added to the reaction (and products are continuously removed)
and a
semi-batch reaction involves the intermittent addition of reactants and/or the
intermittent
removal of products. During such reactions, additional catalyst components
(i.e.
catalyst, MAO, aluminum alkyl) are added to the reactor to sustain the
reaction.
It is within the scope of this invention to optimize catalyst flows after
start up (and
subsequent addition of borate) is completed. Most notably, the relative flow
rates of the
catalyst components might be optimized to maximize rates (e.g. further
chromium and
borate addition) or reduce costs (reduce MAO flow).
Inventive Experiment 5 MAO/Borate
A 600 mL reactor fitted with a stirrer was purged 3 times with argon while
heated
at 80 C. The reactor was then cooled to 45 C (-2 C below reaction temperature)
and a
solution of MAO (0.44 g, 7.1 weight % Al in MMAO-3A solution in isopentane) in
70.36
g of cyclohexane was transferred via a stainless steel cannula to the reactor,
followed
by 55.1 g of chlorobenzene and 19.9 g cyclohexane. The stirrer was started and
set to
1700 rpm. The reactor was then pressurized to 30 bar with ethylene, 10.39 of a
cyclohexane solution of i-PrN(P(2-F-C6H4)2)2 (0.934 mg, 2.68 x10-3 mmol) and
chromium acetylacetonate (1.379 mg, 2.76 x10-3 mmol) was added to the reactor
with
an additional 5 bar ethylene, and the temperature adjusted to 47 C. The
reaction was
allowed to proceed for 15 minutes while maintaining the temperature at 46 C
and the
pressure at 35 bar, at which point Roctadecy1)2MeN1-111B(C6F5)4] (3.820 mg,
3.10 x10-3
mmol) in cyclohexane was added to the reactor. The reaction was allowed to
proceed
for an additional 5 minutes and then terminated by stopping ethylene flow to
the reactor
and cooling the contents to 30 C. Stirring was stopped and reactor slowly
depressurized to atmospheric pressure. Reactor was then opened and product
mixture
transferred to a pre-weighed flask containing 1.5 g of isopropanol. The mass
of product
produced was 100.5 g. A sample of the liquid product was analyzed by GC-FID.
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CA 02758126 2011-11-08
Inventive Experiment 6 MAO/Borate in chlorobenzene
A 600 mL reactor fitted with a stirrer was purged 3 times with argon while
heated
at 80 C. The reactor was then cooled to 45 C (-2 C below reaction temperature)
and a
solution of MAO (0.17 g, 7.1 weight % Al in MMAO-3A solution in isopentane) in
106.23
g of chlorobenzene was transferred via a stainless steel cannula to the
reactor, followed
by 53.4 g of chlorobenzene. The stirrer was started and set to 1732 rpm. The
reactor
was then pressurized to 30 bar with ethylene, 12.0 g of a chlorobenzene
solution of i-
PrN(P(2-F-C6F14)2)2 (0.484 mg, 0.97 x10-3 mmol) and chromium acetylacetonate
(0.327
mg, 0.94 x10-3 mmol) was added to the reactor with an additional 5 bar
ethylene, and
the temperature adjusted to 48 C. The reaction was allowed to proceed for 20.5
minutes while maintaining the temperature at 48 C and the pressure at 35 bar,
at which
point [(octadecy1)2Mehl][B(C6F5)41 (1.367 mg, 1.1 x10-3 mmol) in chlorobenzene
was
added to the reactor. The reaction was allowed to proceed for an additional 9
minutes
and then terminated by stopping ethylene flow to the reactor and cooling the
contents to
30 C. Stirring was stopped and reactor slowly depressurized to atmospheric
pressure.
Reactor was then opened and product mixture transferred to a pre-weighed flask
containing 1.5 g of isopropanol. The mass of product produced was 99.0 g. A
sample
of the liquid product was analyzed by GC-FID.
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TABLE 1
' Runs Al:Cr Cr Total Al Al(MAO) Activity PE
wt% C6 C8 C10 and
Ratio (microM) (microM) (microM) gProductJgCr/hr (based on
(wt%) (wt%) higher
(mol:mol) (based on isolated
(wt%)
isolated product) product)
1 300 41.4 12419 12419 599,011 ' 11.4 18.0 ' 66.8
15.2
2 300 47.3 14187 14187 891,275 2.3 16.6 _ 70.2
13.1 . ci
3 420 5.09 2102 2102 No reaction observed - - -
-
' 4* 442 14.3 6330 6330 1,286,351 0.2 14.9 73.5
11.12 0
"
(1266)
...3
5$ 433 14.9 6463 6463 2,167,639 0.4 19.5 68.4
12.1 C
H
(1267)
tv
66 479 5.97 2855 2855 4,141,822 1.8 22.7 68.4
8.9 . a,
(1270)
ors)
1-,
*1.2 equivalent of [(octadecy1)2MeN1-1][B(C6F5)4Cr was added to reactor 20
minutes into the experiment; i-HeptN(P(C.F15)2)2 ligand used.
$
i
1.2 equivalent of Roctadecyl)2MeNHEB(C6F5)41:Cr was added to reactor 15
minutes into the experiment; i-PrN(P(2-F-C.F14)2)2 ligand used. 55.1 g of
chlorobenzene was used as co-solvent.
1-,
* 1.2 equivalent of Roctadecy1)2MeN1113(C.F5)41Cr was added to reactor 20
minutes into the experiment; i-PrN(P(2-F-C,F14)2)2 ligand used. The experiment
was done in chlorobenzene only.
oi
co
1 0
37
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