CA2193431A1 - Polymerization reactor and process - Google Patents

Abstract

A chemical reactor has an intensively mixed zone and a less intensively mixed zone together with a re-circulation loop to exchange reactants between the two zones. The reactor is suitable for reactions in which it is desirable to add a cold feed stream to the reactor without causing undue precipitation of the warmer reactor contents (and, accordingly, is especially suitable for the preparation of linear low density polyethylene by a solution process). The cold feed is intensively mixed with a portion of the hot reactor contents in the micromixed zone. The intensive mixing is sufficient to present precipitation. The resulting "warm"
mixture is then circulated into the bulk of the hot reactor contents which are located in a reaction zone that is far less intensively mixed.

Description

._ Polymenization Reactor and Process FIELD OF THE INVENTION
The present invention relates to a reactor which is particularly suitable for the solution poly",eri~dtion of olefins.
BACKGROUND OF THE INVENTION
There are a number of processes in which a relatively cooler o stream of liquid is introd~ ~ced into a relatively warmer solution. One of the concer.,s is the precipitation of solute from the warmer solution. One way to minimize this problem is to provide for rapid mixing of the solutions typically using some type of high intensity shear device such as a paddle or agitator stirrer. Generally as the conce"l,alion of solute increases the solution becor"es less Newtonian and the rapid mixing of the relatively cooler solution and warm solution becomes more difficult. The problem is acce"l,Jated if the residence time in the mixer should be relatively short.
Further difficulties arise if the solute is difficult to re-dissolve in the solvent. This may lead to "stranding" or "spaghetti" or precipitate within the mixer which may ultimately affect the product. This problem is particularly acute where the process is consL,ai"ed by "heat balance"
issues.
All of the above issues are particularly relevant to bulk mass and solution pol~ll,eri,alions (as opposed to emulsion and suspension in which the diluent is usually water and heat of reaction is not a significant problem) in which there is a need to remove the heat of polymerization from a reactor. For some reactions this has led to the use of chains of 1 24can.doc 2 2 1 q343 1 reactors with the reactants being heated to successively higher te",peralures and successively higher conversion in dirrerent reactors.
Generally where the residence time in a reactor is relatively long (e.g. in the order of hours) and where the mixing time is relatively short (e.g. in the order of tens of minutes) there may not be too significant a problem.
In the solution poly",eri alion of olefins there are several problems.
The residence time in the reactor is typically quite short and the lifetime of the catalyst at higher te",,l~eralures is also relatively short. Accordingly it is necess~ry to thoroughly mix the catalyst with the solution of alpha olefin "~o"GI~er and polymer quickly. After the catalyst comes to the operaling temperature of the reactor it has a short half life. The situation becomes worse where the viscosity of the solution rises (e.g. high concent,dlion of 2 o polymer or cooler solutions - which may be used to produce higher molecular weight polymer).
Thus it will be apparent that the design of a polymeri,alion reaction especially one used for the solution polymerization of olefins will involve compromises involving a desire to match mixing rates to rates of reaction and balancing such matters as power consumption heat removal capacity and reliability against capital and operali"g costs.
There have been several approaches to this problem. One ap~r~ach has been to use tubular reactors. The high surface area of tube or loop reactors assist in the removal of heat of reaction. However to avoid problems of polymer precipitation the catalyst feed should be at te",peralures above the precipitation temperature of the polymer from the F i ~ 1 24can.doc 3 21q343~

solvent. The obvious answer to this problem is to increase the exit temperature of the tube or loop reactor to increase the temperature profile along the tube or loop at higher ter"perdlures but this i"creased te",peralure typically leads to the formation of more low molecular weight polymer. The problem is accenluated if the polymeri~alion catalyst decays with te",peralure and time. Thus there is a limitation on the temperature increase along the length of a tube or loop reactor. As a result tube reactors tend to be run at relatively lower polymer Col ,cenlrations to avoid problems with precipitate. In order to mitigate these problems tubular reactors also are typically run "~di~h~tically' (i.e. the heat of polymerization GA~ses a te",peral.lre rise between the reactor enlrd"ce and exit with the increase being proportional to the 2 o amount of polymer prod~ ~ced).
Stirred tank reactGr~ used in the production of precipitating polymers under solution conditions may not be limited to using feed stream tel"peralures above the polymer precipitation tel"peralure the same way the tubular reactors are. The mixing in such reactors however is very demanding both on the micro and mauo level. If the micromixing is not sufficient the catalyst efficiency usually suffers beç~ Ise it decol"poses before the ",onoi"er diffuses to it. The ",acro",ixing is particularly important if the polymer produced precipitates at low temperatures. When the temperature of the feed streams into such reactors is below the precipitation temperature of the polymer some polymer precipitation may occur but if the macromixing is very rapid then p~ 124can.doc 4 21 9:3431 the time during which the polymer precipitates is short and thus the amount of precipitated polymer is small (and it will re-dissolve during the time the solution spends in the reactor). If the macromixing is slow cold regions and dead spots form and polymer precipitates in them (thus impeding the mixing which son,eli",es c~uses the reactor to plug with precipitated polymer). The amount of the precipitated polymer and the time it takes to re-dissolve it increases very rapidly as the average reactor temperature decreases. In many cases it is desirable to operale the reactor at low tell,peral-lres so as, for exa",pl_, to facilitate the production of a high molecular weight polymer. The mixing times required in such reactors can be so short that it makes it ir"pra~;tical to design fully mixed reactors which would operale at low temperatures with cold feed streams.

2 o This may still be manageable with low viscosity solutions but with polymersolutions which are very viscous the power requi~e",enls may become unrealistic. Usually a cor"prol"ise must be made for example the feed temperal"re or the operating te",peral,Jre are increased. The former leads to drastically reduced throughput of the polymer the latter to a polymer of comparatively poor properties. The prior art acknowledges these difficulties but doesr, l provide many viable alternatives.
U.S. patent 4 283,339 issued Aug. 4 1981 and assigned to National Distillers and C~,ei"i~l Corp. teaches a process for the solution polymerization of alpha olefins in which dual autoclaves are used in tandem. The first reactor is a relatively higher pressure reactor (e.g. 30 000 psi). The product from the first high pressure reactor is F ~ 1 24can.doc 5 cooled to keep the polymer in solution and avoid precipitation and then introd~ Iced into a second reactor at relatively lower pressures (e.g. 22 000 psi) and the poly",eri~dlion is finished. The referel,ce does not teach or s~ ~ggest the type of mixed reactor of the present invention.
U.S. patent 4 496 698 issued Jan. 29 1985 and assigned to The Dow Chemical Company takes a similar approach to the poly",eri alion of ethylene in which the first reactor is operaled at pressures of greater than 50 000 kPa (about 7 500 psi) and the polyme, i~alion mass is cooled and then fed through a cooling heat exchanger to a second reaclor which may be a tube or loop rea~;tor.
The paper Circulation Time Prediction In The Scale-Up Of Polymerization Reactors With Helical Ribbon Agifators by D.F. Ryan L.P.B.M. Janssen and L.L. van Dierendonck Chemical Engineering Science Vol. 43 No. 8 pp. 1961-1966 1988 discl ~sses a number of devices to stir a tank reactor containing a non-Newtonian solution but does not suggest the micromixing/l"acroi"ixing reactor of the present invention.
Coi ,ce~lually it is highly desirable to achieve micromixing (to ensure good mixing of catalyst and feed) and n,acro,nixing (so as to avoid precipitation at minimal operating cost).
In the reactor of our invention we have essentially separat~d the micromixing and the macromixing functions.
Thus the rea.;tor has a small mixing zone in which the inlet feed is rapidly mixed with reactor coi ,lenls in a desired ratio. A very important , ~ c ~1 24can.doc 6 element of this concept is the recognition that rapid mixing times (i.e. a fraction of the bulk circulation time) can be achieved economically if the blending zone is physically small particularly in highly viscous solutions.
One or more high intensi(y mixing zones with or without the inlet streams being fed into these zones can be included in one reactor.
The mixed reactanls are then transported through the volume of the reactor where the reaction proceeds. After a certain time the reactor CGntel IIS return back to the high intensity mixing zone where they are mixed with the feed stream. The product is preferably withdrawn from the reactor at a region immedi~tely upsl,ea", from the intensive mixing zone.
If ",ecl,a"ical devices are used both in mixing and in inducing the circulation the two devices may either be driven by separale mechanisms 2 o or be driven by the same mechanism such as a rotaling or oscillating shaft.
As already described the reactor conlenls circulate through the reactor volume and the high i"le, Isily mixing zone. It is advanlageous that the distribution of the hold-up times in the re-circ~ ted reactor contents be narrow. That is to say the reactor CGI ,lenls are re-ciru ~-ted in a loop without much axial mixing due to short-circuiting between the forward and return flow in the reactor. This could theoretically be achieved by physically separating the forward and return flows by constructing the reactor as a loop.
Conceptually the mixing device in the high intensity mixing zone is selected so that it operales at the viscosities likely to be encountered 1 24can.doc 7 21 ~3431 during the operation of the reactor. It can be designed to provide the required mixing only or it can also provide or aid in t,anspG, lil ,9 the mixed sl, ear"s through the reactor volume. Desirable mixing of highly viscous reactor co"lents with feed sl,ear"s of low viscosity can be obtained by using a radial impeller designed so that a localized re-circulation zone for rapid mixing and blending is produce-l Mixing can also be achieved by o using the energy of the feed stream in the form of a jet. A simple jet may provide sufficient mixing, or the jet may impinge on a solid surface which may be so sculptured as to provide flow separalion, acceleration or folding similar to those obtained by static mixing devices known in the trade. In the addition to mixing, the jet kinetic energy may also be used to induce the circulation in the reactor, thus a reactor with no moving parts can be designed. The circulation of the reactor contents may also be induced by mechanical devices such as a rolaling screw or helical ribbon in high viscosity envirui ""ents or a combination of rotors and stators in intermediate or low viscosity solutions. Aller"alively other principles can conceptually be used, for example one can imagine that the circulation could be induced by a difference in bulk density between the forward and return flow which is produced by introducing gas to drive an upward flow.
In such a scenario, the drive gas would disengage at the top of the reactor and the liquid would descend. The drive gas might be IIIGi ,ori,er, inert gas, or vapors gei ,eraled by boiling the contents of the reactor.
Thus, conce,clually, it is ir,lpollal,l to achieve both micromixing and macromixing at an ecGnol"ical cost. The prior art appears to recognize r ~ 1 24can.doc 8 -this problem but doesn't provide viable answers. The Applicants have now invented a specific reactor .:lepencle"l process in which these concepls have been put into practical practice.
The invention is particularly suitable for the solution polymeri~ation of olefins especi~lly when the reactor feed is "cold" (below the "critical te""~erdlllre" i.e. below the temperature at which the polymer will prec;pilale from the solution).
To the best of the Applicants' knowledge there is no art disclosing the application of re-circulation to a mixing zone in the solution process for the polymerization of olefin polymers.
The present invention seeks to provide a process in which one or more fresh rea~;tants is mixed in a rapid mixing zone with a portion of reactanls representative of the reactants leaving a rea,;tor and a significant propol lion of the reactanls representative of the reactants leaving the reaction zone are re-circl ~ted through the mixing zone. This process provides an addiliol,al method to help control a reaction as it increases the operator's freedom for example by permitting the operator to control the ratio of the feed rate of fresh reactants and the re-circulation rate of the reactants representative of the product leaving the reactor.
SUMMARY OF THE INVENTION
Accordingly the present invention provides a process for introducing a relatively smaller volume of one or more fresh reactants into a relatively larger volume of one or more reactal ,ls which are representative of the desired output of a reactor then introducing said one 'Q1 24can.doc 9 2 1 ~3431 or more fresh reactants into a portion of said reactants represenldti~e of the desired output of the reactor under conditions of rapid mixing in a rapid mixing zone (or"micromixing zone ) and then introducing the resultant mixture into the bulk of the reactants represenlali~e of the reactor output in a reactor which are pumped through said reactor and a significant portion of which is circul~ted back to the rapid mixing zone and a small portion withdrawn as a product.
DETAILED DESCRIPTION
As noted above the ratio of the re-circulation rate of the reactants representative of the output of the reactor to the rapid mixing zone to the feed rate of said one or more reactanls to the rapid mixing zone may be varied. The rate of re~irculation of reactants represenlalive of the output of the reactor to the feed rate of the fresh reactants to the rapid mixing zone should be greater than 1:1. Preferably the ratio of re-circulation rate of reacta"ls represenlalive of the reactor output to fresh react&nls to the rapid mixing zone is greater than 4:1 preferably greater than 6:1 most pre~erably greater than 8:1. It is desirable to have the ratio high to create as thorough a mixture as practical of reactanls representative of the reactor output and fresh rea.;tants and which will be closer in characteristics to the bulk propei lies of the reactants represenlali-/e of the reactor output.
To a great extent the efficiency of mixing is controlled by the relative viscosity of the components. Typically in accorda"ce with the present invention the ratio of viscosity of the one or more fresh reactanls 124can.doc 10 to the viscosity or the r eacta, lls represenlali~e of the, eactor output is from 10:1 to 1:10 000. Typically the viscosity of the one or more fresh reactan~s will be less than the viscosity of the reaclanls representative of the reactor output (i.e. typically the ratio will be less than 1). Preferably the viscosity of the reactants representative of the output of the reactor is less than 20 Pa s more prererably less than 2 Pa s.
Generally the process of the present invention may be used with an exoll ,er",ic reaction and prererably but not necess~rily the mixture of the one or more fresh reactants and the reactanls representative of the reactor output are cooled before being returned to the mixing zone.
Generally the one or more fresh reaclanls have a te",perature of at least 100~C ~.rererably at least 200~C cooler than the te",perature of the 2 o reactanls representative of the output of the reactor. If the mixture of said one or more fresh reactants with said rea~;tants representative of the rector output is cooled before returning to the mixing zone the cooling may be carried out by passing the mixture through a cooling device such a jacketed tube or pipe or a cooling heat excl ,anger either before or after passing through the quisscel ll zone (e.g. the main reador). If the reaction is highly exothermic it is prererable to maximize the temperature differential between the temperate of the one or more fresh, eactanls and the re~L~, ll3 representative of the reactor output. Additionally the resulting blend may be further cooled before being returned to the quiescent zone and thereby subsequently to the mixing zone. However care should be used to prevent undesired reactions. For example in the p~ 124can.doc 1 1 2 ~ 93~3 1 -case of mixing a stream of relatively cooler and less viscous rea~;la, lls with a warmer more viscous stream of reactanls containing a higher co"cenlralion of product, therel may bel precipitation of the product (e.g. polymer) in the form of stands of "spaghetti". Elimination of "spaghetti" may be accomplished by controlling the ratio of reactants represei ,lali~/e of the reactor output and fresh reactanls to keep the initially blended solution at condiliGi ,s at which precipitation will not occur.
However, care must be taken to avoid or reduce conditions so that a significant amount of catalyst is consumed during initial mixing at too high a temperature to produce a low molecular weight fraction (someli",es referred to as "grease"). While excess production of such a low molecular weight fraction is detrimental, a controlled amount of low ",olec~ ~lar weight 2 o fraction is desir~ble to modify or control certain propei lies of the resulting polymer such as process~hility or heat sealability or even tack in applications such as cling wrap film (or pallet wrap film). The separation of the mixing zone from the reaction zone together with a controlled recycle between the two zones provides the process engineer and polymer chemist with an additional degree of r,eedoi" in producing polymer.
While it is not essential, it is preferred that the zone of rapid mixing and the quiescent or reaction zones be se~,ardled by reactor intervals.
The rapid mixing zone may comprise a smaller volume rea.;tor having a means for rapid mixing such as an impeller, or an impeller in cci"b.nalion with one or more static mixers. The mixing may be accomplished by a jet F~ 124can.doc 12 21~3431 -vortex mixer. In cases where the viscosi~ of the stream of reactanls representative of the reactor output and the stream(s) of fresh feed(s) are relatively low the mixing may be accor,l,~lished using static mixing devices.
The re-circulation of reactants through the quiescent or reaction zone and back to the rapid mixing zone is prererably carried out under conditions of low shear. As the fresh rea.;tants and the reactants o representative of the reactor output have been as thoroughly mixed as possible further mixing is not required nor in most circun)slances desired.
Rather there need only be sufficient agitation of the rea~lanls in the reactor zone to move the reactants through the zone and provide for heat exchange (e.g. either healing or cooling) typically by co"lacl with the walls of the rea-;tor or by an immersion heat exchanger. The circulation through the reaction zone may be provided by a screw or ribbon mixer.
The re-circulation of the reactants back to the zone of rapid mixing may be under more severe conditions utilizing such equipment as a jet or centrifugal pump or a progressive cavity pump.
In a particularly preferred embodiment the present invention is used in association with the solution polymerization of olefins to ~.repare linear low density polyethylene. The l"onon,ers typically comprise a solution of at least 60% by weight of one or more C2.3 alpha olefins optionally with up to 20 weight % of one or more C~2 ,urerera~ly C4 8.
olefin ",o.,o",ers preferably alpha olefins which may be straight chained orbra"ched. rlefelledmG"GI"ersareethylene propylene 1-butene 1-l,exel,e and 1-octene. The ",onol"er may be dissolved in up to about - 5 '~. ~ 'Q1 24can.doc 1 3 2~ 93431 . .
40 weight % of an aliphatic solvent typically having from about 6 to 12 preferably from about 6 to 8 carbon atoms at pressures up to about 5000psi.
The catalyst may be a Ziegler Natta type catalyst or a combination of such a catalyst with a vanadium component together with an aluminum co",pound as an activator optionally in the presence of a magnesium compound soluble in the solvent. The catalyst may be a single site catalyst such as the so-called metallocene catalysts which may be activated with an alu",inoxane such as (methyl aluminoxane MAO") or an ionic activator.
Typically the feed streams comprise the catalyst components which may be prepared by in line mixing upstream of the rapid mixing zone and the monomer stream. The feed streams may at a te",peraLure from -20~C
to 1 30~C typically from about 1 5~C to 50~C.
The feed ~l,ea",s may be fed separalely to the rapid mixing zone or a small reactor. There they are mixed for a relatively short period of time (typically less than about 2 preferably less than 1 minute) with a larger propo, lioll of reactants represel ,lali~/e of the product leaving the reactor.
The ratios of mixing and the relative te""~eralures of the stream of one or more fresh reactants and the reactants re~,resenlalive of the product leaving the reactor have been described above.
The thoroughly mixed product is then fed to a larger reaction zone where the reactants react at higher te",perdl,Jre again for a relatively short time typically less than about 15 prererably less than about 8 minutes. A

.; ~. - c.'Q124can.doc 1 4 ~ . ~
significant proportion of the reactanls may be recycled back to the rapid mixing zone, optionally through a cooler and a smaller propGrlion of the reactanls re~resenlalive if the product of the reaction is drawn off and separate.J from the solvent. The conversion of the feed to product in the quiescent zone should be at least about 75%.
BRIEF DESCRIPTION OF THE FIGURE
Further details of the prefe, led embodiment will now be described with rerere"ce to Figure 1 which is a schematic representation of a prefer, ed reactor embodiment according to the invention.
DETAILED DESCRIPTION
Rerer~ i"g to Figure 1, the reactor (schematically represel,led by the dotted line 7) incl~ ~des an intensively mixed, or "micromixed", zone 2 having a volume V, and a co",parati~ely quiescent reaction zone 3 having a volume V2.
The micromixed zone 2 is mixed by an impeller 4 having a diameter D.
A portion of the conle"ls of the reactor is cira ~l ~ted between the micromixed zone 2 and the reaction zone 3 as illusll aled by the arrow 5 which indicates a re-circulation pattern. The volumetric flow rate of this re-circulation loop is QR.
Cold reactants are fed into the intensively mixed (micromixed) zone 2 through cold reactor feedline 1. The volumetric flow rate of cold reactanls is des~ ibed by the term QF.
Hot reactor products exit through product discharge line 6.

F ~ 'Q124can.doc 1 5 For clarity it should be noted that the micromixed zone 2 and the comparatively quiescent reaction zone 3 may be different zones of a single reactor as schematically suggesled in Figure 1 (or altematively the micromixed zone 2 and the quiescent reaction zone 3 may be physically separate reactors).

Thus in the speci~' case of the solution polymerization of olefins o cold reactor feed (containing ",onoi"er and solvent and optionally comonomer) is fed through the reactor feed line 1 into the micromixed zone 2 at a volumetric flow rate of QF(m3/s). Catalyst is most pr~rerably fed through the same reactor feed line 1 (though optionally the use of a sepa,ale catalyst line (not shown) is also a viable option).
The impeller 4 provides intensive mixing so that the cold feed does 2 0 not cause precipitation of the warmer conte, lls of the micromixed zone 2.
A re-circulation flow (in~' ~ted by reference numerals 5) transports reacta"ls between the micromixed zone 2 and the comparalively quiescent reaction zone 3. As indicated by the subscript h (i.e. rt:rerence numeral 5h), hot reactanls enter the micromixed zone 2 so as to allow the intensive mixing with the cold feed. The resulting warm mixture of hot and cold reactants is indiç~led by the subscript w (i.e. rerere"ce numeral 5~,) which enters the con,parali~/ely q~ _scent reaction zone 3.
The co"~pdr~li\/ely quiescent reaction zone 3 is most preferably stirred or mixed or agitated to allow good dispersion of the heat of poly",eri~alion through the reactor contenls. However this reaction psc,~"~ 124can.doc 16 2193~31 .
zone 3 is "comparatively quiesce"l' (i.e. in con,parison to the "inlensively mixed" or micromixed zone 2).
The product solution (i.e. a solution of polyethylene in solvent) exits through product discharge line 6.
The reactor of this invention allows the solution poly",eri~alion of olefins using cold reactor feed at a tei"perature of more than 100~C lower than the polymerization reactor without causing problem precipitation and without applying the large amount of energy which would be required to micromix the entire volume of reactan~s. Thus this rea.;tor allows a very energy efficient polymerization reaction (as a) the use of cold feed can be used as a heat sink to remove heat of poly",eri,alion and b) this desirable result is achieved by "micromixing" only a small volume of the 2 0 reaCtantS).
P~ arer, ad reactor conditions for the solution polymerization of ethylene (optionally with a comG,~oi"er such as butene or octene) are provided in Table 1.

Condition rl erer, ~d Highly Frerer, ed 5<V, +V2~ 180 40~V1 +V2~70 3 ~ QF QF
2 2~QC~50 8~QC~12 QF QF
3 1 <V2<80 8'V2~10 4 4 c Re, ~ 100 000 100 ~ Re, c 500 where:
QF- Feed rate 1 ~91 24Can dOC 1 7 ~1 93431 QC - [Re-C;rC~ tion flow (QR)] plus [feed flow (QF)]
V, - Volume of micromixing zone 2 V2 - Volume of quiescent zone 3 Re, - Impeller Reynolds number (pND2/ll) in micromixing zone N - Revolutions per second of impeller D - Impeller dia",eter p - Density of rea~;tor co"lents in micromixing zone ~L - Absolute viscosity of reactor co"lents in micromixing zone Additional details are provided in the following non-limiting examples.
Example Ethylene was polymerized in a 600 ml reactor which had the internal diameter of 75.5 mm and length of 150 mm. A stirrer shaft was fitted through the centre of the reaclor top. The shaft was coupled to the stirrer drive through a magnetic coupling. There was 1/4 inch outside diameter OD" (about 0.635 cm OD) ll,er",owell reaching from the reactor top down 30 mm off centre. In this thermowell there were inserted two ll,er",ocouples measuring temperature at about 30 mm and 120 mm from the top of the reactor cavity respectively. Ethylene cyclol ,exane solvent and 1-butene were fed into the reactor through a 1/4 inch (0.635 cm) port that was located in the centre of the reactor at its bottom. These three components were mixed before entering the reactor and the temperature of the mixed stream was controlled by heating the cyclohexane stream to obtain the desired temperalure of the mixed stream. The reactor pressure F ~ 'Q124can.doc 18 was 20 Mpa in all cases. Catalyst was injected at the reactor bottom through a port located 30 mm from the reactor centre. The catalyst was pre~,ared by in-line mixing of a 6 mM solution in cyclohexane of TiCI4/VOCI3 20/80 by mole with a 10 mM solution in cyclohexane of triethyl aluminum. Shortly after these catalyst precursor solutions were mixed, the resultant stream was in,ected into the reactor. The feed rate of the transition metals was kept constant and for each run the amount of triethyl aluminum was optimized to obtain the maximum ethylene conversion.
The changes in conversion were measured as the dirrerei ,ce between the temperature of the reactor inlet stream and that of the reactor outlet stream.
Comparative Example (Conventional Reactor) The reactor stirrer shaft was fitted with two impellers. A four bladed axial impeller was located 70 mm from the top of the reaclor cavity. The blades of this impeller were 18 mm wide and were angled at 45 degrees so as to pump downward. The second impeller was a four-bladed radial impeller with blades 12.7 mm wide and it was located 100 mm from the top of the reactor cavity. Both impellers were of 38 mm diameter. The stirrer operated at 1300 RPM. This high rate of agitation provides "micromixing"
(i.e. intensive mixing). The Reynolds number was low (less than 10) due to the small reactor size. (Larger, commercial size reactors typically have a Reynolds number of from 100 to 500.) The reactor feed rate was 60 kg/hour of cyclohexane and 4 kg/hour of ethylene and 3.5 kg/hour of 1-butene into the reactor. The feed temperature was 11 0~C, the outlet 124can.doc 19 temperature was 180~C. The two tl,er",ocouples in the reactor indicated temperature of 165~C at the bottom and 160~C at the top. The feed temperalure was then lowered in small steps at a rate of about 5~C per hour. When the feed te",peralure reached 105~C the insidé temperatures became unstable and the temperature at the reactor bottom was lower than that at the rea~or top. Shortly after that the conversion was lost and the reactor stirrer shaft seized. The reactor was flashed with solvent and opened. There were polymer strands wrapped around the propeller blades the shaft and the thermowell. This clearly indicates that the solid polymer formed at the reactor inlet did not re-dissolve in the warmer part of the reactor and accumulated in the reactor.
Inventive Example 1 The 600 ml reactor described above was divided by means of internal reactor elements into a micromixed zone and a co",parali~/ely quiescent reaction zone.
The micromixed zone 2 had a volume V, of 60 ml and the quiescent zone 3 had a volume V2 of 540 ml (for a VJ V, ratio of 9).
The volumetric flow ratio of [re-circulation flow QR + feed flow QF] -[feed flow QF] was controlled at 6/1.

The residence time obtained by dividing the total volume (V, + V2) by the feed flow QF, was between 15 and 20 seconds.
The Reynolds number was again low (less than 10) due to the small reactor size. It will be appreci~ted by those skilled in the art that a r i , ~ Q124can.doC 20 -larger reactor will require a higher Reynolds number (e.g. from 100 to 500 for the case of a commercially sized ethylene polymerization reactor).
The flow rate was 60 kg/hour of cyclohexane solvent 4 kg/hour of ethylene and 3.5 kg/hour of butene-1. The feed tei"peralure was initially set at 110~C and the reactor discharge te",per~ilure (or outlet te" ,peralure) was 1 80~C. The previously described thermocouples indicated actual te",peralures of 175~C and 160~C. The feed temperature was then d~upped in small increments at a rate of about 5~C per hour.
Stable reactor conditions (no problematic precipitation) were obtained with a cold feed te",peralure of 75~C (in coi"parison to the failure of the conventional r~ador in the comparative example at a feed te",peralure of 1 05~C.

Inventive Example 2 The conditions of this example were similar to those of Example 1 except that the feed rate of ethylene was increased to 6 kg/hour. The reactor was stable (no problematic precipitation) at a feed tei"persilure of 50~C.

~-o' l',: . ~12~can.doc 21

Claims (16)

1. A process for introducing a relatively smaller volume of one or more fresh reactants into a relatively larger volume of one or more reactants which are representative of the desired output of a reactor then introducing said one or more fresh reactants into a portion of said reactants representative of the desired output of the reactor under conditions of rapid mixing in a rapid mixing zone and then introducing the resultant mixture into the bulk of the reactants representative of the reactor output in a reactor which are pumped through said reactor and a significant portion of which is circulated back to said rapid mixing zone and a small portion withdrawn as a product.

2. The process according to claim 1 wherein the ratio of the re-circulation rate of the reactants representative of the output of the reactor to the feed rate of said one or more fresh reactants is from 4:1 to 12:1.

3. The process according to claim 1 wherein the viscosity of the reactants representative of the output of the reactor is less than 20 Pa s.

4. The process according to claim 1 wherein the reaction is exothermic.

5. The process according to claim 1 wherein said resultant mixture is cooled before being returned to the mixing zone.

6. The process according to claim 1 wherein said mixing zone is in a separate reactor.

7. The process according to claim 1 wherein said rapid mixing is performed by an impeller.

8. The process according to claim 1 wherein said rapid mixing is performed by a jet vortex inlet.

9. The process according to claim 1 wherein said relatively warmer stream comprises a solution of at least 60 weight % of a mixture of at least 80 weight % of ethylene and up to 20 weight % of one or more monomers selected from the group consisting of C4-12 copolymerizable alpha olefin monomers in up to 40 weight % of a C6-12 aliphatic or aromatic solvent.

10. The process according to claim 9 wherein the conversion of said relatively warmer stream on leaving said reactor is not less than 75%.

11. A process for the solution polymerization of at least one olefin monomer in solvent characterized in that said process is undertaken by providing said monomer to at least one reactor having:

(a) an intensively mixed micromixed zone;
(b) a comparatively quiescent reaction zone;
(c) a feed line; and (d) a discharge line for discharging a polymer in solvent solution;
with the further provisos that:
(e) said feed line is directed to said micromixed zone and (f) a portion of said olefin monomer and solvent are circulated from said micromixed zone to said comparatively quiescent reaction zone and back to said micromixed zone at a volumetric circulation rate QR.

12. The process according to claim 11 wherein said at least one olefin monomer consists of ethylene and optionally a comonomer selected from butene and octene and wherein:
(a) said micromixed zone has a volume V1;
(b) said comparatively quiescent reaction zone has a volume V2;
(c) said solvent and said at least one olefin monomer are initially introduced through said feed line at a volumetric flow rate QF; and (d) said circulation rate plus said flow rate equal to a combined volumetric flow rate QC;
with the provisos that:

; and

13. The process according to claim 12 with the further proviso that:

14. The process according to claim 12 wherein the temperature of said flow rate QF in said feed line is at least 50°C lower than the temperature of said comparatively quiescent reaction zone.

15. The process according to claim 14 wherein said micromixed zone is intensively mixed by an impeller having a Reynolds number of from 4 to 10,000.

16. The process according to claim 15 wherein said impeller number is from 100 to 500.