Fast Liquid-phase Processes In Turbulent Flows

FASTLIQUID-PHASE IN PROCESSES TURBULENTFLOWS

K.S. Minskert, AI.AI. Berlin’, V.F! Zakharov2and G.E. Zaikov3

’

N.N. Semenov’s Institute of Chemical Physics, Russian Academy of Sciences, Moscow, Russia ([email protected])

* Bashkir State University, Ufa, Russia ([email protected]) N.M. Emanuel’s Institute of Biochemical Physics, Russian Academy of Sciences, Moscow, Russia ([email protected])

Translated by E. Yu.Kharitonova

/I/ vsPI/// Utrecht

Boston, 2004

VSP an imprint of Brill Academic Publishers P.O. Box 346 3700 AH Zeist The Netherlands

Tel: +31 30 692 5790 Fax: +31 30 693 2081 [email protected] www.bril1.nl www.vsppub.com

0Copyright 2004 by Koninklijke Brill NV, Leiden, The Netherlands. Koninklijke Brill NV incorporates the imprints Brill Academic Publishers, Martinus Nijhoff Publishers and VSP.

First published in 2004 ISBN 90-6764-409-9

All rights reserved. No part of this publication may be reproduced, stored in a retrieval system, or transmitted in any form or by any means, electronic, mechanical, photocopying, recording or otherwise, without the prior permission of the copyright owner. A C.I.P. record for this book is available from the Library of Congress

Printed in The Netherlands by Ridderprint bv, Ridderkerk.

2

CONTENTS INTRODUCTION

5

CHAPTER 1. MACROKINETICS OF FAST CHEMICAL REACTIONS 1.1. The classification of chemical technology reactors 1.1.1. Plug-flow reactors Stirring reactors 1.1.2. Hydrodynamic structure of reaction mkture movement 1.1.3. Problems of instrumental typography under fast chemical processes 1.1.4. Realization of fast chemical reactions in tubular regime 1.2. 1.2.1. Fundamental regularities of fast chemical reactions in turbulent floH ’S 1.2.1.1. Existence o f several macroscopic regimes 1.2.1.2. Correlation between geometry o f reaction zone with kinetic and hvdrodvnamic uarameters 1.2.1.3. The influence o f reagents movement linear rate on uolvmer’s molecular-masses characteristics and vield 1.2.1.4. Effectivenesso f external heat removal 1.2.2. Realization of fast processes in tubular turbulent apparatus 1.3. Intensification of heat- and mass-transfer processes in turbulent regime 1.3.1. Turbulent mixing of one-phase mediums 1.3.2. Turbulent mixing of multi-phase mediums Regulation of thermal conditions underfast chemical reactions in 1.3.3. turbulentflows 1.3.3.1. Adiabatic regime 1.3.3.2. Internal heat removal 1.3.3.3. External heat removal 1.3.3.4. Intensification o f convection heat-transfer in tubular auuaratus 1.3.4. Hydrodynamic and heat similarity criteria and theirfunction in heat- and mass-transfer processes 1.4. Diffuse limitations under some polymers synthesis 1.4.1. Stereospecific polymerization of isoprene on Ziegler-Natta catalysts 1.4.2. Pentadiene-1,3 cationic polymerization Ethylene and propylene copolymerization 1.4.3. Chlorination of butyl rubber 1.4.4.

7 7

CHAPTER 2. 2.1. 2.1.1. 2.1.2. 2.1.3. 2.1.4. 2.1.5.

35 35 35 35 36 36

2.1.6. 2.1.7. 2.1.8. 2.1.9. 2.2. 2.2.1.

EXPERIMENTAL PART Experimental work Oligomerization of piperylene Isoprene stereospecific polymerization Chlorination of butyl rubber Ethylene and propylene copolymerization Formation of macro-structures of reaction and mixing fronts in turbulent flo ws Suspensions preparation Emulsions preparation Studying of hydro-dynamic structure of reaction medium motion in tubular turbulent apparatus Investigation of convection heat-transfer Analysis methods Determination of polymers microstructure

7

8 9 11 12 12 12 13 14 14 15 16 16 19 21 21 22 23 24 24 26 26 29 31 32

37 38 39 40 43 43 43

3

2.2.2. 2.2.3. 2.2.4. 2.2.5. 2.2.6.

Determination of polymers molecular-masses characteristics Determination of chlorine content in rubber Determination of double bonds content in polymers Determination of density and viscosity of liquid flows Analysis of suspension dispersed structure

DEVELOPMENT OF SCIENTIFICALLY-GROUNDED APPROACH TO THE SELECTION OF OPTIMAL REACTION ZONE GEOMETRY UNDER FAST PROCESSES 3.1. One-phase systems 3.1.1. Mathematical modelling 3.1.2. Turbulent mixing in tubular apparatus Distribution of reagents residence times in apparatus of cylindrical 3.1.3. and divergent-convergent types Fast processes proceeding in high-viscosity mediums 3.1.4. 3.1.4.1. Self-similar regime o f reaction medium flow 3.1.4.2. Optimization o f characteristics o f reagents turbulent mixing in selfsimilar regime 3.2. Multi-phase systems 3.2.1. Mathematical modelling The apparatus geometry effect 3.2.2. The reagents introduction method effect 3.2.3. Separation of reaction medium under multi-phase systemsflow in 3.2.4. tubular turbulent apparatus

44 44 44 44 44

CHAPTER 3.

EXTERNAL HEAT REMOVAL UNDER FAST CHEMICAL REACTIONS IN TURBULENT REGIME Quasi-plug-flow mode in turbulent flows as a key to realization of 4.1. fast chemical processes in quasi-isothermal conditions 4.1.1. The influence of fast chemical reaction kinetic parameters on conditions of quasi-plug-flow mode formation in turbulent flows The influence of liquidjhwsphysical charactehtics on conditions of 4.1.2. quasi-plug-flow mode formation 4.2. Regulation of thermal conditions under fast chemical reactions in turbulent regime 4.2.1. Selection of thermal conditions control mode under fast liquid-phase reactions in turbulent regime on the account of external heat removal 4.2.2. Intensification of convention heat-transfer 4.2.2.1. Apparatus geometrv effect 4.2.2.2. The correlation between heat-transfer coefficient and hvdro-dvnamic regime o f reaction medium motion

46 46 46 49 54 63 63 66 74 74 80 82 85

CHAPTER 4.

CHAPTER 5. 5.1. 5.2. 5.2.1. 5.2.2. 5.3.

NOVEL SOLUTIONS IN THE FIELD OF FAST CHEMICAL PROCESSES REALIZATION IN TURBULENT REGIME UNDER POLYMERS SYNTHESIS Butyl rubber halogenation Cationic piperylene oligomerization Kinetics of piperylene cationic oligomerization The influence of kinetic parameters on polymerization reactor selection Separation of fast and slow stages under polymers synthesis in the presence of Ziegler-Natta catalysts

89 89 89 90 94 96 104 104 110

115 115 116 117 120 123

4

5.3.1. 5.3.2. 5.4. CHAPTER 6. 6.1. 6.2. 6.2.1. 6.2.2. 6.2.3. 6.2.4. 6.3. 6.3.1. 6.3.2. 6.3.3.

Isoprene stereospecific polymerization Ethylene and propylene copolymerization Fast chemical reactions followed by solid phase formation

124 132 136

DEVELOPMENT OF NOVEL TECHNOLOGIES FOR FAST PROCESSES IN TURBULENT REGIME Plug-flow tubular turbulent apparatus - the novel type of industrial apparatus Polymers production Ethylene-propylene rubbers production Chlorobutyl rubber production Liquid oligo-piperylene rubber production Anti-agglomeratorproductionfor synthetic rubbers Low-molecular products production in the flow Unblended benzene production Alkyl-sulfates production Neutralization of acid and alkali mediums

141 143 143 147 149 151 151 151 152 158

CONCLUSION

161

REFERENCES

164

141

5

INTRODUCTION Various processes in which reagents in consequence of interaction undergo deep conversions followed by changing of internal structure, substances composition and aggregate state are realized in chemistry and applied chemistry. Chemical-engineering processes use to contain the physical ones also, in particular liquids movement, their mixing, heating and cooling, dispersion, etc. It is worth saying that technology of various chemical products production as a rule comprises chemical and physical processes of the same type characterized by common regularities. But for all that processes of different productions are carried out in apparatus of analogous principles of operation. According to the classification of applying chemistry apparatus there are only two theoretical (limit) models of instruments: apparatus of ideal displacement and of ideal mixing. Real ever-functioning reactors are representing apparatus of intermediate type with some approximation to apparatus of mixing or displacement. It is because particles residence time in them distributes more uniformly than in ideal displacement apparatus but never becomes level as in ideal mixing apparatus. Intensification of chemical processes, reducing of their steel intensity, energy- and resource economy, creation of scientifically well-founded, economical, compact technologies of high ecological safety and apparatus for them are the priority directions of modem technology development. It is obvious that for such problem-solving routine investigations of processes and apparatus in production quantities and conditions are usually complex and prolonged. That is why modelling is very significant; it allows studying regularities of processes proceeding and apparatus functioning with the help of models under conditions permitting application of resulting data independently from production quantities. It is possible to assume that kinetic and hydrodynamic methods of reactors analysis and design are well advanced at present. Methods of computer simulation and modelling are widely used. So, we can say that if we know processes kinetic and hydrodynamic parameters and fundamental particularities of reactor functioning we can calculate all process characteristics and its structure, we also can predict effectiveness of apparatus operation and consumer properties of chemical production. Meanwhile criterion function development for calculation and organization of novel processes and optimization of present productions require overcoming of a big number of problems in production practice. This principle is satisfactory enough for processes with low or medium reactions rates, when creation of isothermal conditions in apparatus is easy. In this case it is easy to calculate and reproduce in working conditions all characteristics of chemical process and to control the last ones. We may see the inverse picture under fast processes proceeding with high rates characteristic duration of chemical reaction is less than 0,001-0,l seconds. This class of reactions was not studied well enough until recent times, as it turned out that it comprised a large number of important chemical and mass-transfer physical processes realized in industry. When carrying out chemical reactions proceeding with high rates (characteristic time of reaction is less than time of transfer) complex temperature and concentration fields are formed in the reaction zone. Fast chemical processes proceed in diffusion field as a rule and they are limited by reagents inlet into reaction zone. So, special approach to the technological scheme apparatus design is needed in every concrete situation. Modern productions processes with application of standard stirred tank reactors are ill-controlled, proceed under conditions wide of the isothermal. They also characterized by unwarranted long times of reagents residence in reaction zone, noticeable reduction of final product total yield, high content of side substances, deteriorated heat removal, etc. When one synthesizes polymers in standard stirred tank reactors undesirable change of their molecular structure, molecular mass (MM), molecular-mass distribution, formation of heterogeneous products and consequently deterioration are observed. Uncontrolled reaction temperature regime may lead to sudden effects, in particular strong hydro- and pneumatic strokes, rejections, heat explosions, etc. That is why selected for fast processes in liquid phase and widely used in industry tank reactors of continuous action (V = 2 + 30 m3) with intensive mixing and developed heat removing system never can be found successful and optimal.

6

The use of ideal plug-flow reactors of continuous action for the same aims is even grater unsuccessful because even under relatively slow chemical reactions realization of isothermal regime in reactors of such type is impossible. Plug-flow reactors are usually function in adiabatic or intermediate but not isothermal (including external heat removal) regimes. So, we may conclude, that practically all reactors using in industry for fast processes are non-optimally adjusted and generally ineffective in production. Consequently the quality of resulted polymers are not optimal and processes are engineering, socially and economically imperfect (reduction of resulted product yield, deterioration of its quality, high content of non utilizable waste, unjustified high consumption of raw materials, electric power, etc.). Thus, under realization of quite fast chemical processes in real productions it is necessary to take into account that true chemical mechanism of proceeding processes is complicated by totality of physical processes such us diffusion, heat- and mass-transfer, hydrodynamics. Fast processes proceed in diffusion region and require macro-kinetic approach to a solution of this problem. For the first time deep investigations of ultra-speed liquid-phase processes began to develop by the example of isobutylene oligomerization and polymerization as a classical model reaction. Sufficiently great number of fast chemical and mass-transfer physical processes was studied then. Necessity of decrease of reaction apparatus outline dimensions down to sizes commensurable with “torch” sizes proved to be a very important scientific and applied conclusion. This lead to the development of compact highly productive tubular turbulent apparatus functioning in a novel quasiideal plug-flow regime. Novel technology of fast chemical reactions realization in turbulent regime is also effective under many mass-transfer processes (extraction, mixing, mixing of flows with various viscosity and density, dispersion, saturation of liquid mediums by gases, etc.). Application of tubular turbulent apparatus allows increasing effectiveness of external heat removal that simultaneously determines possibility of their use as a compact heat exchangers. Accordingly to aforesaid fundamental regularities of fast liquid-phase chemical, heat- and masstransfer processes in turbulent flows are considered in the book. Novel energy- and resource-saving technologies on the base of compact tubular turbulent apparatus are presented.

CHAPTER 1. MACROKINETICS OF FAST CHEMICAL REACTIONS

1.1 The classification of chemical technology reactors Reactors take central place in every production connected with chemical conversion of substances and serve as key elements of technological scheme. Reliability and stability of realized process, its manufacturability, economy and environmental safety depend on perfection of reactors [ 1, 21. Chemical reactors applied in industry differ in forms; it causes troubles with their classification. In particular, it was supposed before that there was no correlation between type of realized chemical reaction and conditions of equipment operation in which this reaction was carried out, but this is not the case. Idealized models are usually applied for theoretical design and selection of reactor constructions with consideration of proceeding polymerization process kinetics [3-61. Degree of reliability of one or another model is determined by the depth of investigation both chemical kinetics and hydrodynamics and thermal physics of reaction apparatus. Such investigations are prolonged for years very often. That is why they have to put real industrial process in operation directing attention to empirical description of process rough-kinetics and the simplest models of temperatures and concentrations distributions, such as continuous or periodic reactors of ideal mixing and plug-flow or their empirical combinations.

1.1.1.Plug-jlow reactors Tubular and columnar apparatus (apparatus length-to-diameter ratio U d > 100) including screw equipment relate to plug-flow reactors type [7, 81. Plug-flow reactors are applied for many of gas-phase reactions realized in production quantities, in particular for ethylene polymerization under high pressure conditions [9], and for some liquid-phase reactions, for example polystyrene synthesis in columns and other rubbers and plastics productions. Near 10% of polymer and 30% of fibers manufacture are produced in apparatus of such types [ 101. Model of ideal desaturation (model with plug flow regime) is the favorable approximation for calculation of reactor parameters [3,4,6]: any cross-section normal for flow, weight hour space velocity w and flow’s properties (pressure, temperature and reaction mixture structure) are uniform; narrow distribution of reagents residence times in reaction zone T ~ diffusion ~ ; in the axial line (coplanar mixing or turbulence) in comparison with weight hour space velocity is negligible low. If the polymerization process is of the second order (proposing that active sites concentration is not changed in the course of process) for the plug-flow reactor monomer concentration is changed according to the equation [8]:

1)

C M = C MO exf-k P Cas w

(1.1)

where C,O, C,, C,, - concentrations of monomer initial and at the reactor’s output, and active sites concentration; w - discharge of reaction mixture; vr - volume of reaction zone (reactor’s volume); k, - constant of chain propagation reaction rate. We can receive correlation between reaction zone sizes vr and monomer conversion degree x out of ( 1 . 1 ) :

Equation (1.2) allows to calculate the optimal reaction zone size in plug-flow reactor necessary for achieving of required monomer conversion degree in dependence on kinetic parameters of chemical process k, and needed monomer conversion degree x.

8

If there is a laminar flow as it is in the majority of polymerization plug-flow reactors (especially under block polymerization [lo]) cross sizes, in particular plug-flow reactor diameter dd, are often limited by transport processes rate in cross direction. Transport processes rates have molecular character and consequently low value. At the same time long-term industrial processes (characteristic chemical reaction time is high zch) lead to the necessity of long plug-flow reactors use and creation of considerable motive powers. For polymerization processes in solution or block (bulk) when polymer is dissolved in monomer the effect is intensified by increasing viscosity of reaction mass especially in the cases when movement is stimulated by external pressure as in flow reactors. It is impossible to realize isothermal regime in plug-flow reactors because variation of heat emission along reaction zone length accordingly to heat generation kinetics is required for that. That is why plug-flow reactors usually work in adiabatic or at least non-isothermal regime (with external heat removal) [ 113. The equation of thermal balance for stationary regime in micro-volume of plugflow reactor is as the following [ 101: dT dP Q = - C -dV+q-dv -2KnRdL(Tad -Tx), (1.3) P dz dz P where q - reaction heat effect; P - product yield per volume unit; IC - coefficient of heat transfer through the wall; Tad- adiabatic rise of temperature in reaction zone. Any deviations from ideal displacement regime are called stirring or inverse stirring. In that case tank reactors of ideal stirring are the idealized model of continuously operating apparatus in opposition to ideal plug-flow reactors.

1.1.2. Stirring reactors Vertical and horizontal capacitive apparatus appointed by mechanical mixers of impeller, turbine, ribbon, disk, extrusion type including stirring rod for high viscosity mediums mixing are referred to stirring reactors or cascade of flow reactors [3-5, 101. However, application of apparatus with mechanical mixing for polymers synthesis is limited by maximum limit of viscosity 10-lo2 Pa.sec. Nevertheless, about 90% of plastics and synthetic rubbers, and 70% of fibers are produced in stirred tank reactors [lo]. The model of ideal mixing for tank reactors formulation is based on the following [ 3 , 6 , 121: -the full mixing is reached in apparatus and this fact allows to utilize reaction space to the maximum without noticeable formation of stagnant zones; - structure of reaction mixture at the reactor outlet and in its volume differs from reaction mixture composition at the inlet; - the distribution of times of reagents residence in apparatus is broad; - formation of isothermal regime is possible, by this temperature field can be varied in the course of the process (in a stirring reactors cascade as a rule). Thermal regime in stirred reactors is characterized by adiabatic rise of the temperature in the reaction zone [ 1, 81:

where p, Cr - density and thermal capacity of reaction mixture accordingly, and by the rate of heat removal (without boiling consideration) [ 101: (1.5) Q , =wpCr(Tr -To)+KF(Tr -Tc) here Tr, Tc - temperatures in reactor and of coolant accordingly, F - surface of heat exchange. Calculation of continuously running of stirred reactor under polymerization is based on the solution of mass balance equation by monomer [8]:

9

dC

J= WCMO- WCMl

-v

v

r P Here V, = kC,C, - polymerization rate of the second order in approximation to equality of active sites and catalyst concentrations C,. Transforming (1.6) with consideration of monomer conversion x = (C,o-C,)/C,o one can receive the expression for calculation of reaction zone sizes in stirred tank reactor of ideal mixing:

dz

Calculations on the base of (1.2) and (1.7) showed [8, 131 that when productivity of stirred reactor was equal to that of plug-flow reactor reaction volume of the last one was always lower. By this, the higher degree of reagents conversion and total order of reaction, the bigger the difference between stirred and plug-flow reactors volumes. Consequently plug-flow reactors are more favorable by chemical process kinetic parameters than stirred reactors, especially when volumes vr are equal. However stirred reactors in practice are more comfortable during their use and economically sound (because of the simplicity of their manufacture). Idealized models of mixing and displacement that are determined by hydrodynamic structure of liquid flows are usually used for theoretical calculation and selection of reactor construction accordingly to the kinetics of process.

1.1.3. Hydrodynamic structure of reaction mixture movement One can receive the most perfect information about hydrodynamic structure of reagents movement in reaction zone if flow instantaneous speed in any point of apparatus is known. This speed defines temperatures and reagents concentrations fields. As a consequence, the hydrodynamic structure of reagents flow in reaction zone significantly effects the rate of chemical-engineering processes and its registration is the integral part of industrial apparatus modelling under realization of concrete chemical process [3,7]. The measuring of flow speed in any point of industrial apparatus is very complicated, and sometimes practically impossible. So, one can determine hydrodynamic structure of reaction mixture movement indirectly, in particular by studying of fluid particles distribution by their residence times in reaction zone. For measuring of aleatory variable, i.e. particle residence time in apparatus, this particle should be marked in some way allowing to register its entering and coming out of apparatus and to receive concentration curve of flow at the outlet. This curve is called output curve or response curve [ 14, 151. One can hardly recognize the mechanism of processes occurring in apparatus using such approach because real speeds field in it is not known. That is in this case apparatus is considered as “black box”. At the same time such analysis method of flow structure in reaction zone is easy enough. Quantitative data processing is essentially simplified because they definite the function of only one variable - time. Furthermore, the data of residence times distribution guarantee reliable estimation of real flow structure in apparatus, Le. allow to “look into” “black box” [3, 7, 161. Distribution of reagents residence times in reaction zone is described by dependence [ 171:

Fi4 =

c.1(4

00

(1.8)

Here Ci - concentration of reagents with residence time zi in reaction zone. Experimentally received response curves are convenient to present by derived distribution curve by residence time in non-dimensional coordinates C - 0 that are calculated out of ratios [18] C = CivdCo and O=zi/i, where CO- amount of introducing indicator, = vdw - calculation residence

10

time of indicator determining as apparatus volume Vr to volumetric speed of reaction mixture movement w ratio. Cellular and diffusion models are usually used for estimation of longitudinal mixing (turbulence) in reaction zone and consequently for evaluation of deviation degree of fluids hydrodynamic structure from ideal displacement and mixing regimes [3,7, 19-21]. Accordingly to cellular model fluids structure and response curve shape in apparatus are described by distribution differential function: nnOn- le- ra0 N(0)=

(n-1)!

Here n - the only parameter of cellular model equal to cells (reactors) number in cascade of ideal stirred reactors, ideal stirring regime is achieved at n + [3]. It is accepted [22], that if cells number in reactor n 2 8, then such apparatus can be calculated as plug-flow reactor with enough for industrial practice accuracy. Accordingly to diffusion model any deviation in reagents residence times distribution from the distribution at ideal mixing independently of reasons is a consequence of longitudinal mixing and is determined by turbulence level [7, 16, 231. Diffusion model means piston flow with superimposed effects conditioned by molecular diffusion, presence of small curls, stagnant zones and radial gradients of fluids speeds. In this case fictitious diffusion coefficient E (coefficient of longitudinal mixing) is used [17]. And mono-directional diffusion (along x axis of apparatus) is described by the correlation [3]: OQ

dC dC -=-w-++dz dx

d2C (1.10) d g Solution of equation (1.10) at impact introducing of indicator leads to expression for function of residence time distribution [17,20,21,24]: (1.11)

L2 w where Bo = - - Bodenstain’s criterion (or Pekle’s criterion for longitudinal mixing PeL). Flow vrE structure is judged by numerical value of criterion Bo comparing its quantitative deviation from ideal displacement (at Bo + w) and ideal mixing (at Bo + 0) regimes. Comparison of experimental derived curves of distribution by reagents residence times with curves calculated from (1.9) and (1.11) allows to receive numerical values of Bo H n and consequently to estimate degree of deviation of flows structures in reaction zone of various geometry from idealized models. Relative residence time 0, is a very important parameter describing the character of reagents movement in reaction zone. For the estimation of relative reagents residence time in reaction zone by experimental response curves medium-integral magnitude is zr calculated [22]: W

I zfi?(ddz

z = i

’

o 1 w

(1.12)

I Ci(dd.t 0 In this case relative residence time of fluids in reaction zone characterizing stagnant zones and reagents slips (bypassing) is determined by ratio 0,= 2, / i. If 0,< 1 then there is a slip of reagents in reaction zone that is undesirable under real processes of polymer production. When 0,=

11

-

1 medium-integral magnitude of actual residence time zn corresponds to calculated z and hydrodynamic regime in reaction zone approximates to regime of ideal displacement [23]. Thus, the differences between stirred and plug-flow reactors are determined by hydrodynamic structure of reaction mixture movement. The last one is the main criterion for scientifically based selection of reaction zone geometry with consideration of specificity of occurring process. However the regularities of fast chemical processes proceeding with characteristic reaction time zx e 0,Ol sec condition significant complications when choosing the optimal geometry for reaction zone. This is due to the fact that basic theses of ideal mixing and displacement models under local fast processes do not work [ 11.

1.1.4.Problems of instrumental typography under fast chemical processes In chemistry there are examples of extremely fast processes proceeding in liquid phase, when chemical reactions rates are comparable or higher than rates of reagents' mixing. A lot of processes of ion and free-radical polymerization, chlorination, hydrochlorination and sulfation of olefins, alkylation of alkanes by alkenes, chlorination of aromatic hydrocarbons, etc. relate to such extremely fast processes [ 1, 251. Cationic (electrophilic) polymerization of isobutylene that may be considered as a classical model of fast chemical reaction attracts the greatest attention for theoretical investigation among fast polymerization reactions [26, 271. It is because this process is theoretically clear in general and of high practical importance. In industry oligo- and polyisobutylenes with molecular mass (MM) 112-50000 and higher are produced by cationic polymerization of isobutylene in the presence of AlC13 in hydrocarbons (butane and others) or chlorinated hydrocarbons (ethyl or methyl chloride, etc.) medium at 173-353 K in stirred reactors of 1,5130 m3 volume and complex construction [26]. Reactors are equipped by advanced inside and outside heat-exchanging surfaces (130 m2 and higher) with liquid ethylene or ammonia as refrigerants, intensive mixers providing linear rates of reaction mass movement of 1110 d s e c order for cooling and necessary productivity control according to thermal and material balance. Medium reagents residence time in reaction zone is about (1,8+3,6).103 sec [l]. Constants of initiation and chain propagation (ki, kp)are lower than 105+106Vmole.sec [27341, and determine characteristic time of chemical reaction equal to 10-'110'3 sec at real initial catalyst and monomer concentrations. It means that isobutylene polymerization proceeds by 100% in places where reagents are introduced into the apparatus (within less than 11-10cm from catalyst introduction) and process is limited by reagents mixing, i.e. process proceeds in diffusion field. Heat accumulation in the reaction system at high turbulence of reaction mixture movement leads to the oscillation temperature change in reaction zone in time [28]. In topochemical aspect fast polymerization reaction proceeds as "torch" with characteristic gradients of reagents temperature and concentration by coordinates of reaction zone [28, 351. Shape of forming reaction front is analogous to that under burning reaction [36, 371. In this case the main reactor volume is not enabled in chemical process and presents ballast in a sense increasing reagents residence time in reaction zone that promotes by-products formation. Furthermore, reaction zone doesn't reach reactor's walls and the use of external heat removal is ineffective [ l , 2, 25, 271. That is why thermostating of fast polymerization processes in industrial and laboratory production is difficult enough. Gradients of polymers M M and MMD are observed in process course in reaction zone volume due to non-isothermality [38, 391. It is important that dependence of medium M M on reagents concentration, and also significant MMD broadening at initial monomer concentration increase are the macroscopic consequences of local polymerization proceeding. At the same time polymer M M according to kinetic scheme [28] is totally determined by chain transfer to monomer and predicts independence of M M and MMD on catalyst and monomer concentration in isothermal regime [l]. Moreover, catalyst and monomer solutions are introduced into reactor irregularly as a rule and because of high polymerization rates catalyst and monomer don't have enough time to mix well with reaction mass.

12

This particularity causes additional non-stability in reactor operating, MMD broadening and reduction of resulting product MM [28]. It was found that mentioned regularities of fast polymerization were applied to any fast processes including low-molecular synthesis in flow [31,40-441. Thus, using traditional approach for instrumental typography of fast chemical reactions both under polymer synthesis and low-molecular products production in flow leads to the uncontrolled processes and nonoptimal quality of resulted products [ l , 25,26,31,39,45].That is why it is of great importance to find optimal conditions for fast chemical reactions carrying out. It is worth saying that fast chemical processes should be studied and carried out in stream conditions in tubular reaction vessel with ideal mixing, i.e. in high-turbulence flows [ l , 361. It is expedient on account of several reasons. First, carrying out of fast reactions in a stream limited by impermeable wall presents the most convenient and natural way of their investigation. Second, under realization of such reactions in classical stirred tank reactors it is in principle impossible to provide effective heat- and mass-transfer for the time commensurable with chemical reaction time. Third, turbulence conditions of moving stream allow carrying out fast chemical processes with maximum intensity for least time and in minimum volume. Forth, investigation of reactions in stationary stream with turbulence mixing is in principle easier and more effective than under single proceeding under closed volume conditions. At last, under fast reactions realization in stream it is possible to vary characteristic mixing time z,lk = R2/D, (R - radius of stream (flow) limited by impermeable wall) and, in particular to make it commensurable with chemical reaction characteristic time Z& = l(d (k,~constant of active sites deactivation). First of all, it was necessary to solve at least three problems [2]: a) fast mixing of two and more reagents, especially different in density and viscosity (following condition must be fulfilled in ideal - zdXI zch); b) reagents residence time in reaction zone ( T =~ LN)shouldn’t be significantly higher than time of chemical reaction because of slower secondary reactions following condition must be fulfilled in ideal - zr = zch); c) removal of large quantity of heat evolving in substantially small reaction zone volume or realization of conditions when a very fast temperature increase in this zone doesn’t influence on chemical reaction course and quality of resulted polymers. As fast chemical processes proceed in diffusion regime, so for revealing of regularities and specificity of their realization it is necessary to use macro-kinetic approach describing chemical reactions with consideration of mass- and heat-transfer [46, 471. On the base of this approach to modelling of fast polymerization processes regularities of proceeding of such reactions were revealed by the example of isobutylene cationic polymerization [28-31, 38, 39, 481. In particular, it was found that many of problems under fast chemical reactions may be solved by creation of turbulence high enough in reaction zone. The last fact led to the development and wide use of tubular turbulent apparatus in industry [ l , 2, 11, 25,45,47,49,50]. Thus, specificity of fast reactions of polymerization and low-molecular products synthesis in flow determines inadequateness of idealized models both under tank stirred and plug-flow reactors use with classical understanding of their operation (see 1.1.1 and 1.1.2) that caused significant complications under their design. Consequently, traditional technological equipment applied in the majority of productions for fast chemical processes of polymer synthesis are ineffective. In this case compact and highly productive tubular turbulent apparatus of original construction working only in quasi-plug-flow mode is necessary to use.

1.2. Realization of fast chemical reactions in tubular regime

1.2.1. Fundamental regularities of fast chemical reactions in turbulent flows

15

1.2.1.1. Existence o f several macroscouic regimes New phenomenon was found by the example of fast liquid-phase isobutylene electrophylic polymerization - the influence of geometric sizes of reaction zone on resulting products yield, their molecular-mass characteristics and monomer conversion [28, 35, 511. One make out two macroscopic types of process in topochemical aspect: A (plant front) and B (torch). At low radiuses R (type A) reagents mixing is effective and active sites, monomer and temperature are comparatively uniformly distributed by reaction zone coordinates in the form of plane perpendicular to reactor axis (reaction plant front). This fact well provides high (up to 100%) monomer conversion and formation of novel steel unknown quasi-plug-flow mode in highly turbulent flows [52]. Consideration of longitudinal mixing (by the expense of turbulence) distinguishes this regime from plug-flow regime (see 1.1.1) [3-81. “Dithering” of temperature in reaction zone due to longitudinal turbulent diffusion and fluctuations leads to almost the whole reaction volume is at constant temperature (quasi-isothermal regime). As a consequence, fast chemical reactions in production under realization of quasi-thermal regime become easily controlled. The other extreme case (local torch regime) is realized at relatively high values of R (type B) (tank stirred reactor). Active sites deactivates having no time for diffusion into reaction volume peripheral parts that in this case, are the slip zones of non-reacted monomer. As a consequence specific, complex in configuration fields of monomer, active sites concentrations and temperature are formed. Reaction doesn’t reach reactor’s walls and product yield due to monomer slipping is always lower than 100% [38,39].

1.2.1.2. Correlation between geometry of reaction zone with kinetic and hvdrodvnamic parameters Critical radius Rcr determining transformation from torch regime of reactor operating (macroscopic regime B) into plug-flow regime in turbulent flows (type A) depends on the correlation between monomer diffusion (Dt) and active sites deactivation (h) [ l ,28,42-441:

R

R

=IT

= A

(under polymerization)

(1.13)

(under low-molecular products synthesis in flow), because conditions of cr k[C]” approaching of quasi-plug-flow mode are determined by z,,li, = Rz/Dt I Zch = lk (l/k[C]”~’).Three various hydrodynamic regimes of reaction mixture movement: laminar (low R), transition and turbulent were determined under R changing (at constant speed of reaction mixture movement V) in dependence on characteristic time of mixing Znix= R2/Dt 11, 2,471. Under laminar regime diffusion coefficient is low (D = IO-’ m2/sec) and time of mixing is high (it is increasing as Rz).At transition regime effectiveness of mixing rises a little and consequently Znlix reduces. Under turbulent regime Dt is high (> 10” m2/sec) and it increases according to the low Dt R. Thus, there are limitations by minimum and maximum acceptable radius of reaction zone at concrete flow rates V. When R > R,,,, and reagents movement regime is turbulent there is transformation from plan into torch reaction front [51]. At R < Rnin transformation to the transition and then to the laminar regime with sharp reagents mixing deterioration occurs that doesn’t present practical interest. On the whole there is interval of reaction zone radiuses values narrow enough (in dependence on T ~ I V ~ , and consequently Dt) that provides turbulence flow with formation of reaction plan front causing quasi-isothermal regime in reaction zone (Rdn < R < Rm,).

-

14

Correlation between geometric sizes of reaction zone and kinetic parameters of process proceeding in turbulent regime determines possibility of experimental estimation of fast chemical reaction kinetic constants as it was shown in [51-551. Thus, kinetic parameters of polymerization fast processes (kp, b)and linear speed of flow V determine geometric sizes (R, L) and optimal configuration of reaction zone. New possibilities and methods of processes control allowing to regulate monomer conversion and molecular characteristics of resulting polymers, in particular by forced change (limitation) of reaction zone geometric parameters were revealed. 1.2.1.3. The influence o f reagents movement linear rate on volvmer’s molecular-masses characteristics and yield Values of Dt may be varied in wide range at the expense of preliminary turbulization, by modification of mixing method, direction and speed of reagents flow [56, 571. If tubular reactor works at R > Kr and conversion naturally doesn’t achieve 100% then increasing of Dt in several times, in particular at the expense of reagents movement speed V rise leads to increase of conversion level p in spite of (and it is unusual for classical ideas) the fact that reagents residence time reduces (length of reaction zone L is constant, and V increases) [58]. There is graduation of temperature maximums in reaction zone at rising of V and consequently Dt in spite of total polymer yield increase. Temperature dithering (quasi-isothermal regime) leads to the increasing of medium MM and narrowing of MMD of resulting product [ l , 58, 591. 1.2.1.4. Effectiveness o f external heat removal

One of the most important factors determining effective course of chemical process and its manufacturability is temperature field in reaction zone that depends on kinetic and thermodynamic parameters, reagents concentration, speed of reaction mixture movement, turbulence level of flow, apparatus geometry, etc. Analysis of thermal regime under fast polymerization reactions in turbulence regime showed that it is necessary to use internal heat removal (boiling of reaction mass) or its combination with preliminary cooldown of initial crude (autothermal regime) [60, 611. In dependence on heat efficiency of process q and reaction product yield AF’ temperature rise ATad in apparatus may come to hundreds of degrees, and all heat evolves quickly (for seconds or their parts) and at a very small distance along the reactor length Lch = V’Tch = v/~[cI”” (under polymerization LCh = v&) [40,41]. Formation of quasi-plug-flow mode in turbulent flows under fast chemical processes when reaction zone reaches heat exchanging reactor walls determines possibility of application of effective external heat removal that allows regulating of molecular characteristics of resulting polymer [62-651. At external heat removal (Tch = const) reaction mixture temperature change AT=T,d-Tr along cooling zone length Lcwl from maximum Tad = To+ATad (To - initial temperature of reaction mixture) up to required temperature Tr is determined as following [ 1,40,41]: (1.14) Out of this the length Lcwlrequired for maintenance of necessary temperature in reaction zone Tr:

(1.15) Modification of catalyst feeding into reaction zone in particular realization of its distributed introduction (“zone” model of process carrying out) is good possibility for control of polymer MM and MMD [l]. “Zone” model is series connection of several independent reaction zones in which

15

reaction mixture comes after process completion in previous zone and where new portion of catalyst (or monomer) solution is introduced [62, 651. Thus, totality of received results allows to considerate fast liquid-phase processes as independent reaction class that is differed by individual specificity and investigations methodology. The necessity of realization of fast polymerization processes in tubular turbulent apparatus is determined by their basic operating principles: - reduction of apparatus dimensions down to sizes comparable with "torch"; - stream regime of operation that minimizes reagents residence time in reaction zone and consequently possibility of side processes; - possibility of application of external heat removal. Furthermore, tubular turbulent apparatus allow guaranteed and easily control of earlier uncontrollable fast polymerization processes, in particular by the change of [49, 661: a) reaction temperature field; b) coefficient of mass- and heat-transfer; c) turbulence level in reaction zone; d) reagents concentration; e) geometry and sizes of reaction zone; f) way of reagents introduction; g) linear flow speed; h) chemical nature of the solvent; i) pressure in the system; j ) kinetic parameters of chemical reaction. Mentioned regularities of fast chemical reactions in turbulent regime determine wide application of tubular turbulent apparatus in industrial production [ 11.

1.2.2. Realization of fast processes in tubular turbulent apparatus Total profitability of chemical productions in many respects is determined by equipment cost and service and repair bill of main units including reactors forming the central unit of chemicaltechnological scheme [67]. For the reduction of such costs it is necessary to use apparatus of lesser volume that are able to combine high productivity, simplicity of manufacturing and servicing and operation safety that is very problematically when working with high-viscosity solutions in polymer production. The last fact causes the topicality of development of high-performance compact devices and tubular turbulent apparatus are the most manufacturable among them [ l , 2,50, 68, 691. Tubular turbulent apparatus are widely used in various processes of applied chemistry under both polymer and low-molecular products production: oligomers and polyisobutylene mediummolecular sorts [70-741, medium-molecular sorts of butyl rubber [ 11, polyisoprene [75-771, statistical copolymers of ethylene and propylene [78, 791, butadiene-a-methylstyrene rubber [ 801, block-copolymers of ethylene and propylene [ 8 11, ethylene chloride [68], 1,2-dichloroethane [ 11. Tubular turbulent apparatus are also used for synthesis of alkylbenzines by reaction of sulfuric acid alkylation of isoparaffin by olefins [82-871, azo-dyes by interaction of sulfuric acid suspension of organic amine with water solution of sodium nitrite [88-901, for purification of oil and gas condensates from hydrogen sulfide and mercaptans [91-961, blend benzenes [97], etc.. They also may be used in production of piperylene [ l , 981, chlorobutyl rubber [99, 1001, alkyl sulfates [ l ] and others oligomers including neutralization of acid and alkaline mediums [ 1021. Moreover, possibility of creation of high turbulent flows in mixing zone allows effective application of tubular turbulent apparatus as mixers [69], dispersers (for preparation of thin homogenious emulsions) [ 103-1081, extractors [109], absorbers [ 1101, heat exchangers [ 1111, etc.. Realization of fast processes in tubular turbulent apparatus determines [ 11 universality of processes, simplification of technology, reduction of production space, high products quality, decreasing of coefficients of expenses, decreasing of emissions, discharges and yields of illconditioned products, energy saving, metal saving, possibility of ill-conditioned products use,

16

increase of overhaul life of equipment and reactor’s specific output, increase of environmental safety, high social and economic efficiency. Totality of received results allows to considerate fast liquid-phase chemical processes as independent reaction class that is notable for its own investigations methodology, individual specificity and instrumental typography. Burning processes that were marked out as a separate class of oxidation reactions due to characteristic macroscopic features are the close example (but not analogue). Novel class of reactions in practical aspect should be realized in quasi-plug-flow mode in turbulent flows. Furthermore, high level of turbulent mixing in tubular turbulent apparatus allows wide using of them for intensification of heat- and mass-transfer processes. Consequently, it is important enough to study operating efficiency of apparatus of novel generation in dependence on construction parameters and specificity of fast chemical and mass-transfer physical processes under polymer synthesis. It is also important to estimate the degree of deviation of tubular turbulent apparatus operation from idealized mixing and overshooting models with the aim of selection of optimal reaction zone geometry while improving traditional instrumental typography (stirred and plug-flow reactors) for realization of fast processes of polymers synthesis.

1.3. Intensification of heat- and mass-transfer processes in turbulent regime Mixing in liquid phases is the most commonly used method for increasing of turbulence level in reaction zone, intensification of heat- and mass-transfer processes in chemical and other branches of industry [112-1141. With the aim of reception of homogeneous by structure and properties polymer product it is necessary to provide under polymerization the equal conditions in total reaction volume. It is possible to realize by intensive mixing of reaction medium when homogeneous fields of reactants concentrations are formed at zmix < zch condition. Traditionally used for this aim apparatus of large volume with mechanical mixers as a rule can not adequately provide the homogeneity of fields of reactants concentrations along reaction volume under fast chemical reactions especially when working with highviscosity mediums [ 114, 1151. For example, system viscosity under polymerization (from 10” to 10’- lo3 Paesec), high constants of propagation reaction rate (kp> 102.1/mole.sec), a large amount of evolving heat (60-100 kDj/mole) determine sharp rise of temperature in reaction volume at the catalyst and monomer delivery point right up to heat explosion [7]. The interest in compact apparatus allowing along with specific output creating of high turbulence level in mixing zone grew recently [116-1181. As they showed in [119, 1201 intensity of turbulence approaches 50+70% in cylindrical apparatus with turbulence promoters (for undisturbed flow in plain canals this value was l+3%). Apparatus with turbulence promoters [121-1241, including those of divergent-convergent design [125] don’t contain internal moving elements and are able to provide homogeneity of conditions for both chemical and mass-transfer physical processes. Application of tubular turbulent apparatus for processes limited by mass-transfer stage is determined by possibility of their control by increasing of turbulence level under reagents mixing. The last fact required the development of theoretical methods for studying of characteristics of liquid flows turbulence mixing in tubular canals as a base of fast chemical processes carrying out in optimal conditions and intensification of heat- and mass-transfer.

1.3.1. Turbulent mixing of one-phase mediums It is well known that radial introduction of reagents into apparatus in comparison with coaxial method provides better mixing in reaction zone. In particular, in [29, 56, 57, 62, 1261 on the base of Navie-Stocks equations in combination with K-E turbulence model it was shown that significant improvement of flows mixing under other equal conditions can be reached only by change of way of reagents introduction. Change of initial turbulence level influences on mixing characteristics as at coaxial, so at radial introduction. Experimental and calculation data [37, 127, 1281 indicate on

17

essential improvement of mixing and increase of effective mass and heat diffusion if there are circulating zones. Formation of the last ones is reached by application of various mechanisms [ 1291 in combination with the rise of flows speeds [56, 1261. In particular central one-shot catalyst introduction into reaction zone under fast polymerization processes is less effective than catalyst introduction by outside ring. The best results are observed not for coaxial, but for radial introduction of one of the components if there is conical extension in the initial part of reaction zone [56, 591. Nevertheless, it was shown in [ 103, 105, 1151 that mixing of liquid flows differing in density and especially in viscosity is a very complex problem. Suffice it to say that in spite of unlimited solubility of concentrated sulfuric acid (p = 1,8 g/cm3, p = 27,8 mPa9sec) or glycerin (p = 1,26 g/cm3, p = 1490 mPa.sec) with water (p = 1,00 g/cm3, p = 1 mPasec) prolonged conservation of phase interface is observed in tank apparatus even under mechanical mixing. Significant increase of mixing effectiveness of two or more reagents differing in density and viscosity is reached at the expense of rise and stabilization on definite level Dt along reaction zone, in particular when local hydrodynamic resistances of divergent-convergent design are used [ 1,69, 1251. This determines the expediency of investigation of canal geometry influence on turbulent mixing efficiency in tubular apparatus. One should know turbulence characteristics to estimate the characteristic time of reagents mixing. They proposed a lot of various methods for estimation of mixing time in turbulent flows particularly the review on this theme and mixing models classification are presented in [130]. Since in [125] liquid movement is described by average over Reynolds Navie-Stocks equations with the use of K-Eclosing, so the characteristic time of turbulence mixing can be estimated as [ l , 32, 1251: Ttwb = lZmt (1.16) here 1 is characteristic linear size of region where homogeneous reagents concentrations field should be created. If one supposes that turbulence diffusion coefficient Dt is equal to kinematic coefficient of turbulence viscosity vt which in its turn can be expressed by specific kinetic turbulence energy K (m2/sec2)and rate of dissipation of last one is E (m2/sec3)then (1.16) will be as following: (1.17) Tturb = 11,IE ?/ KZ However, correlations (1.16) and (1.17) don’t take into account the fact that in turbulent flow mass transfer processes are carried out at the expense of molecular diffusion and viscosity flows especially under polymer synthesis in a volume small enough [ 1311. That is why if micromixing processes exactly are limitative as under polymerization then it is necessary to use other estimating expressions. In this case, as it was shown in [125] the engulfment model is used frequently enough [132, 1331, according to which characteristic time of micromixing is estimated as following: ~ n l =i 17,3(~/~)O” ~ ~ ~ (1.18) here v is kinematic viscosity. In a number of cases homogenization of medium is limited by exchange processes between large turbulent flows and presenting in them smaller flows, i.e. mesomixing [132, 1341: (1.19) T,,SO = 1+2(12/ E ) 113 Comparison of characteristic mixing times mentioned above allows revealing of limitative mechanism of reagents concentrations field equalization under fast processes carrying out in turbulent regime. For this aim out of values calculating from (1.16)-(1.19) the biggest should be chosen and then it should be compared with chemical reaction characteristic time Tch. If the last one is turned out to be significantly higher, so chemical transformation process occurs in kinetic region and diffusion limitations don’t affect the resulting product structure. If we consider the mixer for medium homogenization without chemical transformation, so limitative mixing time should be compared with average reagents residence time in reaction zone 2,. The fact that in highly turbulent flows liquid viscosity doesn’t influence on main volume medium movement is interesting and important enough [3, 1351. In this case they say that tlow is self-simila in relation to viscosity and the influence of the last one is displayed in a narrow enough wall layer. The value of Reynolds’ criterion above which the self-simila field is observed in many respects is determined by flow geometry. For example, in [3] they showed that under sphere flow the

18

self-simila field in which resistance coefficient doesn't depend on Re and consequently on viscosity comes at Recr= 500. It is very important to study the possibility and conditions of self-simila flow formation in relation to viscosity in reaction zone of cylindrical and divergent-convergent designs that will widen exploitation potential of novel technology under polymer synthesis. The fundamental approaches to definition of turbulent flows macro-kinetics and macro-mixing processes are considered in [136-1391. Special attention was focused on micro-mixing models in the context of method based on equation for density of random variables probabilities distribution. Advantage of this method is that we can calculate average rate of chemical reaction if know the corresponding density of concentration and temperature possibility distribution. Since intensive mixing of liquid flows is observed at developed turbulent regime, so according to [136-1391 one can't directly apply ordinary equations of continuum movement (for example, Navie-Stocks' equation with molecular coefficient of viscosity) in this case. That is why it is generally accepted to use the so-called average equations of turbulent continuum movement, they are also called as average by Reinolds. Numerical solution of equation of continuum turbulent movement in combination with expressions of IC-& turbulence model allows to obtain over the whole tubular apparatus volume the fields of axial and radial speeds, pressure, specific kinetic energy of turbulence, its dissipation and of some other characteristics that are expressed by listed values. This fact permits to calculate the optimal in respect to mixing effectiveness geometry of reaction zone under fast chemical processes including polymerizations in turbulent regime. The most attractive and requiring detailed studying problem is to reveal the conditions of formation in tubular turbulent apparatus of novel unmatched quasi-plug-flow mode in turbulent flows at the expense of creation of corresponding hydrodynamic conditions in mixing zone and purposeful variation of apparatus construction [ 1,28,42-441. Formation of various macroscopic reaction fronts including quasi-plug-flow mode as it turned to be afterwards is the general phenomena and applies to any fast chemical processes [42-441. Moreover, characteristic macrostructures in turbulent flows limiting by impermeable wall can be formed also under mixing of liquid flows without chemical reaction [44]. Particularly, in topochemical aspect, in general case of two liquid flows mixing the formation of five principally different macrostructures of mixing fronts and reaction were experimentally observed: 1. "torch" type - the macrostructure characterized by broadening of reaction front limits as moving off the point of reagents introduction along flow axis; 2. "plan front" of reaction type - perpendicular macrostructure or at some angle of apparatus axis; 3. "drift" type, when process proceeds in the whole reaction volume at the expense of intense inverse diffusion of flows; 4. "wave" type - structure of the wave shape that extends from introduction point of main flow to mixing zone along apparatus axis without cross diffusion; 5 . "rope" type - macrostructure in the form of dual mirror "waves" structures with intensive axial rotation. Formation of characteristic fronts macrostructures under liquid flows mixing and fast chemical reactions including polymerization is mainly determined by the ratio of linear speeds of reagents introduction V1 I V2 (VI - speed of axial flow, V2 - speed of radial flow) and method of their introduction into reaction zone [ 1,42,44]. Macrostructure of "plan front" is optimal (dominant) under fast chemical processes ("knowhow") and corresponds to the quasi-plug-flow mode in turbulent flows. This structure of liquid flows mixing provides effective heat- and mass-exchange under not only fast chemical (polymerization, low-molecular products synthesis), but also mass-transfer physical processes (mixing, dispersion, extraction, etc.). Analysis of conditions of creation of the most important for production quasi-plug-flow mode and mathematical modelling of mixing process of two reacting liquids in the framework of IC-& turbulence model have shown [3 11, that quasi-plug-flow mode formation in turbulent flows under fast chemical reactions is reached at the expense of production of initial components mixture and/or

19

reaction product out of reaction field into reagents introduction zone. Such product recovery is realized by stationary curl resulting from significant difference between flows speeds (pressure gradient). At the same time we think that we can't ignore the possibility of quasi-plug-flow mode formation in turbulent flows under fast chemical reactions when they realized by "excision" of the torch's central part cross sizes of which are significantly larger than reaction zone diameter [ l , 441. Thus, fast polymerization processes should be realized in essentially novel unmatched quasithermal plug-flow regime in turbulent flows. The last one can not be reached neither in mixing reactors, nor in plug-flow apparatus [ l l ] , but can be realized in tubular turbulent apparatus. Possibility of formation of plan reaction front under fast processes of polymer synthesis determines expediency of studying of the influence of reaction rate constant on conditions of quasi-plug-flow mode formation in turbulent flows. Furthermore, there are no literature data on the influence of physical parameters of liquid flows on the formation of characteristic macrostructures of reaction and mixing fronts. This causes definite complications under fast chemical processes realization in optimal regime. Under polymer synthesis there is considerable number of processes proceeding under phase separation conditions, in particular in such systems like "liquid-liquid'' (high-viscosity polymer solution - low-molecular fluid), "gas-liguid" (polymer-analogous processes) and "solid-liquid'' (catalytic polymerization). Necessity of substance transfer from one phase to another by massexchange [ 1401, i.e. removal of diffusive limitations is the feature of heterogeneous processes. In this case one have to deal with serious problems while realizing fast chemical reactions [141]. 1.3.2. Turbulent mixing of multi-phase mediums Problem of creation of multi-phase reaction systems with developed surface of phase contact is especially actual under polymer synthesis. In particular at the stages of reaction mixture formation under emulsion [ 1, 801 and suspension [ 1421 copolymerization, halogenation of elastomers [55, 1431, decomposition and removal of electrophilic catalysts and Ziegler-Natta catalytic systems out of polymer [l], saturation of solvent by monomers [78, 791, formation of heterogeneous and micro-heterogeneous Ziegler-Natta catalytic systems [ 1441 and so on. Surface-volumetric diameter of disperse phase's particles (drops, bubbles, particles) d32 is widely used for description of dispersion processes [18, 107, 108, 121, 1451: 3

d

=-

z midi

(1.20)

32 m i d ; ' It concerned with specific surface area F of spherical particles of disperse phase by correlation: (1.21) here mi - the number of drops of i-fraction with diameter di. Polydispersity coefficient h is one more significant value characterizing the quality of resulting disperse systems [ 1461: z m i d3 i .zmidi (1.22) k = z m i d4i z m i For monodisperse systems h = 1, for polydisperse - h < 1. The less the value of h parameter, the higher particles sizes spread. Among numerous apparatus applied for preparation of disperse systems with developed surfaces of phases contact tank apparatus are the widespread ones. They are characterized by external

20

source of turbulence (mixer) and consequently by developed turbulent flows (intensity of turbulent mixing reaches 30+50%) [147-1491. In [150, 1511 the following ways of reducing of dispersion degree of emulsions formed in apparatus with mixers are proposed: increase of mixing time, volume of crushing zone or local energy dissipation in crushing zone. For reducing of emulsions dispersion degree in apparatus with mechanical mixing the following variant is the most preferable: increase in crushing zone of local value of energy dissipation maximum of which is localized in agitator zone [148]. Moreover, the major factor influencing on hydrodynamic situation in reaction zone is vertical circulation (diffusion in axial direction) [152, 1531. In spite of apparatus with mechanical mixing, rotary-pulsating [118, 154, 1551 and hydroacoustic devices [156-1581 are widely used for realization of processes proceeding on phase interface. Such apparatus are based on intensive cavitation field creation in reagents zone mixing. It is worth saying that these apparatus are characterized by high effectiveness under heterogeneous processes, but they are power-consuming and complex in manufactoring. In [103-105, 107, 108, 110, 1591 it was shown that formation of disperse systems in turbulent flow in tubular apparatus is effective method of fast heterogeneous chemical reactions realization. It was experimentally revealed that high-dispersive emulsions are formed over L = (20+50)&, required energy is provided only by flow speed that should be about 5-15 d s e c [log]. However in [lo51 it was shown that effectiveness of emulsification can be increased by the increase of conical extension angle (>30") at the beginning of reaction zone. At the same time for effective dispersion of flows limited by impermeable wall and differing in density and viscosity Dt equal to 0,065-0,080 m2/sec is necessary along the whole mixing zone extension. This value should be maintained constant over the whole apparatus length [103, 1051. The question about the influence of tubular canal geometry and flows introduction way on dispersion effectiveness in turbulent flows remains unclear. But it can reveal additional parameters allowing to control dispersion degree of reaction mixture and consequently of limitations value for reagents diffusion. It is theoretically possible to estimate minimal size of drop (bubble) &r undergone deformation in turbulent regime from the following correlation characterizing stability of phase interface [ 1471: (1.23) Where i s - surface tension; v' - pulsating speed of element's scale determined accordingly to Kolmagorov-Obukhov's "two thirds law" [ 1601: V' = ( ~ d z ) " ~ (1.24) Pulsations of less scale possess significantly less energy and are not able to deform particles of disperse phase. Pulsations of big scale carry the elements of disperse phase and do not deform their surface. The fundamental problem under estimation of disperse inclusions of multiphase systems in tubular turbulent apparatus according to (1.23) is calculation of rate of turbulence kinetic energy dissipation E. It requires the development of model describing disperse processes in turbulent flows. For numerical studying of multiphase flows regularities including calculation of rate of turbulence specific kinetic energy dissipation the following turbulence models are widely used: standard (KEMODL) [161-1631; modified (KECHEN) [ 1641 and renormalized (KERNG) [165-1671 K-E models. In spite of the fact that two-parameter K-E turbulence model over several years is widely used in calculations of hydrodynamic and heat-mass-exchange in complex turbulent flows some remarks were advanced about it recently 1168, 1691. Calculations made by PHOENICS software [ 170, 17 11 confirm that standard K-E turbulence model really gave too high values in particular of turbulence kinetic energy K. Better results are received when KECHEN and KERNG models are used especially developed for separation flows in canals of complex form, Le. they can be used for calculation of divergent-convergent apparatus [ 1721. Calculations of disperse systems movement with the use of POENZCS software [170, 1711 by modified and renormalized K-Eturbulence models allow obtaining reliable results on turbulence flow characteristics distribution along apparatus volume. Numerical solution of differential equations in partial derivatives with the help of final

LI

volume method allows the receive over reaction zone volume of divergent-convergent design the fields of following functions: axial u and radial v speeds for each phase, pressure p, inclusion volume fractions of continuous a1 and disperse a2 phases, specific kinetic turbulence energy K, its dissipation E and some other characteristics. It should be noted that in the majority of works on multi-phase flows quantitative estimation of disperse inclusions are given. However analytic formulas for calculation of diameter of forming in turbulent flow drops and bubbles are found rarely. Such studies for tubular turbulent apparatus of divergent-convergent design were not carried out at all and this limited designing of this apparatus for reactions proceeding on phases interface under polymer synthesis. Another picture is observed under flow of system "solid-liquid". Under mass-exchange in disperse systems "solid-liquid'' it is necessary to take into account "adhesion effect" - the equality of speed on phase interface to zero and absence of deformation of disperse phase elements under disperse phase influence [173]. At the same time in many cases there is a problem of reception of homogeneous fine suspensions, in particular under high active heterogeneous and microheterogeneous Ziegler-Natta catalytic systems formation for initiation of olefins and dienes polymerization [ 144, 174, 1751, anti-agglomeratorsfor synthetic rubbers [ 176-1781, pigments for dyes and so on. Moreover, increase of phase contact surface leads to significant acceleration of heterogeneous chemical reactions [ 1571. Processes of suspension preparation by condensation method under chemical reaction with sedimentation are widely used in chemical engineering. Chemical reactions underlying the preparation of fine suspensions are ion-exchanging and consequently proceed with high rates [ 1791831. In [183] the correlation between mixing on various levels (macro-, meso- and micro-mixing) and crystallization process in apparatus of periodical action with mixer was revealed. The influence of mixing intensity, reagents concentrations and their volume ratio on particles sizes under chemical reactions with sedimentation was studied. When mixing speed increases at reaction moment reduction of particle sizes and increase of their number occur. However under prepared suspensions processing (it was found on the example of production of yttrium oxalate [181]) speed of mixing practically doesn't influence on particles sizes. It was shown in [179] that up to mixer's rotation speed of 400 rpm increase of particles sizes of disperse phase occurs and further rise of rotation speed (>400 rpm) leads to the decrease of sediment particles radiuses. Authors explain the reduction of specific surface under mixing intensity increase (up to 400 rpm) by the fact that in this case the diffusion rise (feeding of reagents to the crystal surface) is dominate over micro-mixing speed rise [ 1841 leading to the increase of new phase nucleuses number. Thus, increase of intensity of turbulent mixing allows production of fine systems "solid-liquid". Furthermore, reactions underlying the condensate method of their reception proceed fast. This determines expediency of studying of tubular turbulent apparatus application possibility in particular divergent-convergent design for preparation of homogeneous fine suspensions by condensate method under fast chemical reaction.

1.3.3. Regulation of thermal conditions underfast chemical reactions in turbulentflows 1.3.3.1. Adiabatic regime

Temperature field in reaction zone is the major factor determining effectiveness of fast polymerization processes. This field depends on reactor-polymerizer geometry, monomer and catalyst concentrations, speeds of flow movement and turbulence, way of components mixing, etc. [ l , 26,62, 66, 1851. Under polymerization fast processes when R < R,, [28, 351, when intensive longitudinal and cross mixing average temperature in reaction zone so that MMD and average MM are turned to be close to characteristic for isothermal conditions at temperatures corresponding to adiabatic mixture initial heating. However, polymerization possibilities in adiabatic conditions are usually limited by

22

the fact that at high temperatures MM is essentially reduced. At this secondary reactions proceed in particular destruction and macromolecules bonding. Possibility of initial mixture cooling in the case of process under which reaction mixture boiling temperature is not reached is limited by boiling temperature of coolant, for example ethylene (Tbil. = 183 K). Until solvent boiling begins mixture temperature will rise in proportion to polymer yield increase in accordance to the expression [60]:

c

A M = A ( T -T ) (1.25) q s o If we use temperature turn-down equal 90°C (coolant - liquid ethylene), then we shall be able to polymerize about 15 mass % of isobutylene at the expense of adiabatic heating of reaction mixture [60]. If we use liquid ammonia as coolant (Tbil=243 K), then for the time of reaction mixture heating it is possible to receive not more than 10 mass % of polymer. One of the effective method of chemical reaction heat abstraction is internal heat removal under reaction mixture components boiling [27, 61, 62, 186, 1871. 1.3.3.2. Internal heat removal

Boiling occurs in some interval narrow enough, and Tbil depends on amount of polymerized monomer, resulted polymer MM, etc. in a complex way [ l , 261. To a first approximation we may think that boiling occurs at constant temperature especially because for a number of processes evaporation heat is enough to polymerization completion (at monomer initial concentration 13-3,0 mole/l). Reaction mixture temperature rises up to Tbil and then process proceeds at constant temperature until all solvent andor monomer part boil away [60]. When total transformation of boiling liquid into vapor occurs one can observe significant temperature rise in reaction zone due to polymerization of either liquid (if monomer in system doesn't boil) or gas monomer [61]. One can estimate polymer quantity that will be produced at the expense of initial mixture heating from To up to Tbil 1601: (1.26) Out of this expression one can obtain the following:

Here SO= (1-Mo), Mo - the initial mass parts of solvent and monomer in reaction mixture accordingly; %, % - monomer and solvent evaporation heats accordingly. If boiling occurs at variable temperature, then numerical criterion of temperature constancy at polymer MM and MMD analysis is expressed by the following correlation [26]:

Tmax -Tmin mkil (1.28) boil boil << E m - E P where R - gas constant; E,,, E,, - activation energies of chain propagation and chain transfer on the monomer reactions. In special case when constants of chain propagation and active sites deactivation rates do not depend on temperature, polymerization kinetic scheme corresponds to fast initiation and to the first orders of chain propagation reaction on monomer and active sites deactivation on their concentration the expression for polymer yield was received [ I , 621:

k

k

Z Ai . = M

i=l

1

( I .29)

kp i =1

23

Thus, internal heat removal at the expense of reaction mixture boiling is the method of thermostating of fast chemical reactions effective enough. Limitation of reaction mixture temperature under fast polymerization in liquid phase at the expense of monomer or solvent part boiling influences on process proceeding in different ways in dependence on reaction zone radius [61]. In small radiuses field (R < R,,), at internal heat removal when quasi-plug-flow mode in turbulent flows forms and temperature in reaction zone is homogeneously distributed along radius R process course and molecular-mass characteristics of resulting polymer cease to depend on initial reagents temperature in a definite temperatures interval [ 11. Furthermore, in this case reaction zone reaches reaction walls and it is possible to apply external heat removal. 1.3.3.3.External heat removal Particularities of polymerization fast processes in turbulent flows (see 1.1.4) assumed that generally accepted methods of reaction heat removal in process course in particular by external thermostating will differently influence on process macrokinetics and consequently on molecularmass characteristics of resulting polymer products [63, 641. Effective method of thermal regime control is considered in [ 1881, where authors suggest to use reactors-catalytic heat exchangers, when heat-generating and heat-exchanging surfaces are not divided in space. However this method is not acceptable under fast liquid homogeneous processes realization. Analysis of temperature change in reaction flow under fast polymerization shows [ l , 60-651 that the increase of radius reaction zone higher than Rr at the absence of external heat removal leads to significant change of reaction temperature regime. For example at low reaction zone radiuses temperatures of central and peripheral flows differ a little (not more than AT' = 2 4 ° C ) and only at the initial stage of polymerization. Temperature reduction in reaction zone due to heat removal over wall at average conversion degrees (about 50 mass %) leads to decrease of low-molecular fractions amount in resulting polymer [62] If one doesn't take into account the longitudinal heat transfer by turbulent thermal conductivity inside the reactor, then temperature change (cooling) at the expense of external heat removal is determined as following:

In(1-

AT

)9-

-

AT - 2KLcool. pC RV ' Tad - Tch P

(1.30)

Tad -Tch where Tad = ATad + To. It is consequent from (1.30) that heat removal efficiency is in inverse proportion to rector radius at constant linear speed V or in direct proportion at constant volumetric speed w of reaction mixture flow (productivity). That is one should increase R maintaining apparatus productivity approximately constant and trying to improve external heat removal. However at that reagents concentrations and temperature gradients appear (torch regime is formed) otherwise flow regime transforms from turbulent into laminar. Mentioned factors limit the radius of tubular turbulent reactor from the top [ 1,601. Modification of catalyst introduction into reaction zone in particular realization of multi-step catalyst introduction into tubular turbulent apparatus ("zone" model) is the convenient way of polymer MM and MMD control [62]. For "zone" model realization in turbulent regime the following conditions should be fulfilled [ l , 27, 661: 1. reaction zones should not be crossed, Le. the distance between adjacent points of catalyst introduction should be Li > V k ; 2 . the R
d/"d

3.

mode); intensity of heat evolution is not high in each zone, and heat-transfer coefficients should be such to fulfill the requirement of temperature constancy in each reaction zone (quasiisothermal regime).

24

,.

It was showed in [ l , 61-66] that all mentioned conditions are fulfilled at:

pLi VqAM ATi =

Rrf

<
P

'

(1.31)

where ATi - characteristic temperature difference in i reaction zone; PVqAMi - rate of heat evaporation in i reaction zone; &AT& - rate of heat "dithering" along reactor; AT2 - characteristic temperature interval for MM change; Ti - average temperature in i reaction zone. It was shown in [ l , 2, 40, 411 that methods of heat regime control under fast polymerization processes are acceptable also under low-molecular products synthesis in flow. However systematic investigations in the field of searching of regulating ways of heat regime in reaction zone by external heat removal under both polymers and low-molecular products synthesis in turbulent regime were not held. It should be noted that in turbulent flows heat removal at external cooling is determined by heat transfer with the help of heat conductivity and convection, i.e. by convective heat-exchange [3]. As a consequence calculation of heat-transfer coefficient in reaction zone of various geometry is the main problem under fast polymerization processes realization [1891.

1.3.3.4. Intensification o f convection heat-transfer in tubular amaratus Intensification of heat-exchange is one of the essential problems of applied science solution of which allows realizing fast exothermal reactions in optimal conditions. A lot of methods of convective heat-exchange intensification are known at present [ 190- 1931. Turbulization of flows on local hydrodynamic resistances (turbulence promoter) is the most widely used and unique method for one-phase flows. Since canals lengths are usually higher than length of relaxation of flow that was disturbed by turbulence promoter, then to receive noticeable intensification effect the number of turbulence promoters should be big enough [ 1111 and regular. h u r l i n g of tubular apparatus (ratio of broad diameter (divergent) to narrow one (convergent) apparatus section Q/de < 1,l) is widely used now for intensification of wall heat layer [194-1971. Authors suppose that it is not necessary to affect flow core because its turbulization slightly increases heat emission but leads to significant hydrodynamic resistances and energy loses. In accordance to [194, 1971 the highest intensification of heat-exchange is reached in transition region (Re = (2+5).103) and is equal to ap,/ a,,, = 2,83 (heat-transfer coefficients of reaction mixture moving upon profiled and smooth canals, accordingly) at dd / dc = 1,08 and section length L, = (0,25+0,5)dd. the ratio up,/ amis maximum and equal to 2,3 under liquids flow and 3+3,1 under gases movements. Furthermore, the presence of turbulence promoters under profiling of tubular apparatus reduces the rate of obliteration of heat-exchanging canal by scale deposition in 3-5 times [198]. Intensification of convection heat-exchange under deep profiling of tubular canal in the form of divergent-convergent design (dd/dc = 2) was considered in [ l l l , 199, 2001. It was shown that resistance coefficient of tubular turbulent apparatus of divergent-convergent design is approximately 25 times as much in comparison with cylindrical; however it is 15% from calculating sum of local resistances. Calculation of heat-transfer coefficients on the base of criterion models showed high level of intensification of convective heat-exchange in tubular turbulent apparatus of divergent-convergent design, in particular for diapason equal to Re = ( l+2).104 this level turned to be 3,5+4 times [ l l l , 1991. Thus, the application of tubular turbulent apparatus with local hydrodynamic resistances including divergent-convergent design for realization of fast polymerization processes allows additionally to increase of external heat removal at the expense of intensification of convective hestexchange. In [201] the correlation between heat-transfer coefficient and longitudinal mixing coefficient was revealed. This fact determines expediency of revealing of correlation of heat-transfer processes efficiency in tubular turbulent apparatus with hydrodynamic structure of reaction mixture

25

movement by studying of reagents residence times distribution (see 1.1.3). This allows to use the method of response curves for modelling of heat-exchange processes under fast polymerization processes.

1.3.4. Hydrodynamic and heat similarity criteriaand theirfunction in heat- and mass-transferprocesses Under fast polymerization processes as it was mentioned above one usually faces two problems: creation of high turbulence mixing and intensification of heat-exchange. The method of dimensions analysis is often used for studying of chemical apparatus operation regimes and specificity of heat-mass-exchange processes. This method has something in common with similarity theory, widely used for analysis of various physical and physical-chemical phenomena and is the main instrument for approximate description of complex chemical-engineering processes [3, 12, 202, 2031. Lets consider application of methods of similarity theory and dimensions analysis to the hydrodynamic and heat processes of chemical technology with the aim of their possible use for calculation of fast processes under polymer synthesis in turbulent regime. It is well known that fundamental criterion of hydrodynamic similarity is Reynolds' criterion characterizing the influence of frictional force on liquid movement [3, 12, 135, 202,2041:

Re=VdP

P

(1.32)

For similarity of phenomena of reaction mixture forced flow it is enough if it satisfies a numerical claim - equality of Re values. In the case of free liquid flow (natural convection) Granshof criterion is the measure of frictional force-to-lifting force ratio determined by the difference of densities in various points of non-isothermal flow that is the main under laminar flow [3,202,204]:

Gr =

gd%Atp2

(1.33)

P2 where 8 - coefficient of liquid volume expansion; At difference of temperatures between wall and liquid (or vise versa) by the value of which the difference of liquid densities is determined. Damkler's criterion Da is often used for modelling and calculation of reagents mixing under fast polymerization processes [7, 136-1391. The value of Da allows comparing the convection and chemical reaction rates, i.e. the ratio of hydrodynamic and kinetic constituents of any real process. Conformably to fast chemical processes criterion Da may be expressed in the following way (meaning that Da = %x/Tch) [7]: for polymerization processes RLkd Da=-, Dt for low-molecular products synthesis in flow

Da =

R2k[C]"-l

(1.34)

(1.35)

Dt It is obvious that when mixing time is commensurable with characteristic reaction time numerical value is Da = 1. If polymerization reaction is carried out at low monomer conversion Da << 1, and at high conversion Da >> 1. Similarity of heat-transfer processes on interface between reactor wall and moving liquid flow is determined by Nusselt's criterion (Nu) [3, 191-197, 199, 200,2041:

ad NU=-, h where h - medium heat conductivity.

(1.36)

26

For analysis of heat similarity Prandtl's criterion (Pr) is very important. It characterizes similarity of heat-carrying agents' physical properties under convective heat-exchange processes [204]:

Pr=-pcP (1.37) h Criterion Pr is determined by values expressing physical properties of liquid. Applying theory of heat similarity one can calculate heat-transfer coefficient a in dependence on hydrodynamic regime of liquid flows in tubular canals [3]. At laminar regime of reaction mixture flow in tubular canals (Re < 2300) [3]:

Gro9'h -

a = 0,0S5Re0733 R

where px,h,, C,, - viscosity, heat conductivity and heat capacity of flow at wall temperature T,. For transforming regime (Re - 2500-7000) [3,204]:

~ o o 7 @ % ~ ~ % ~ ~ ~ v ~ ~ ti (1.39) RQ1pQ47

a = Q O O ~ Z ~P Q ~~ Q -~ ~ A R

For turbulent regime of liquid flow (Re > 35000) [204]:

Criterion Re is a convenient instrument under studying of reaction mixture movement in apparatus under fast processes including polymerization [ 1, 71, emulsification [ 103, 1051, mixing 11163, etc.. Heat similarity criteria are convenient for investigation of heat-exchange [191, 1921, for calculation of construction and operating regime of chemical reactions [ 1881, creation of effective and compact heat-exchange apparatus [ l l l , 193, 194, 197, 1991, etc.. Moreover, application of tubular apparatus of jet type as a base for creation of modern technology for many of chemical industrial processes requires definite values of Re (Re > (1+1,5).104) in flows mixing zone, i.e. definite hydrodynamic regime of liquids movement. Thus, with the help of criterion equations of heat- and mass-transfer one can model with adequate accuracy the tubular turbulent apparatus operation including conditions of quasi-plug-flow mode formation in turbulent flows. In polymers industrial production with the use of tubular turbulent apparatus for perfection of concrete technological stages limited by mass-exchange one should neatly know specificity of corresponding processes proceeding (heat regime, kinetic parameters of reaction, characteristic reaction and mixing times, etc.). One should also deeply understand effect of turbulence level in apparatus on fast process character as a whole and quality of resulting product.

1.4. Diffuse limitations under some polymers synthesis As it was mentioned above technological realization of polymers synthesis is often complicated by systems viscosity rise in process course and by significant dependence of polymer product quality on temperature and concentration heterogeneity in reaction volume. There is a great number of problems under fast polymerization processes. The main regularities of the last ones were developed by AI.AI. Berlin and K.S. Minsker with collaborators by the example of cationic

27

polymerization of isobutylene in turbulent flows [26-35, 38-44, 47, 49, 51, 52, 56, 59-66, 70-741. Main ideas were confirmed and realized in production of other polymer products in particular copolymers of isobutylene with isoprene (butyl rubber), isobutylene with styrene, divynilstyrol rubber, stereoregular cis-l,4-isoprene rubber and so on [ 11. 1.4.1.Stereospecific polymerization of isoprene on Ziegler-Natta catalysts

Production of high stereoregular synthetic isoprene rubber that is maximum approaching by its properties to natural rubber (NR)is one of the major problems of petrochemical industry. Macromolecules of NR contain up to 98,5-99% of isoprene units of cis-1,4 and the rest of 3,4structure. In industry of Russia technology of cis- 1,4-polyisoprene production is developed and commercialized. In the presence of Ti-AI catalysts "Synthetic rubber SKI-3'' (all-Union State Standard 14925-79) is produced and "Synthetic rubber SKI-5" (technical specifications 2294-05 116810126-96) is received with the help of lanthanide catalysts containing alcohol additive products of chlorides of rare-earth elements of neodymium group [205]. Synthetic trans-1 ,4-polyisoprene (gutta-percha) possesses unique thermo-plastic properties, acid, alkali, oil stable and is characterized by high elastic resistance and is produced on V-AI catalytic systems in industry (content of trans-units is 93-99%) [174, 175,2061. According to classical conceptions [ 174, 1751 stereoregulation under isoprene rubbers synthesis depends on chemical and electron structure of catalytic system' homogeneous component AIR3, structure and molecular parameters of metal (V, Ti, Cr, Nd) salt, etc. [207]. Conditions of active sites (AS) formation play an important role. Structure of units forming at 1,3-dienes polymerization is determined by nature of AS including end unit of chain but isn't determined by monomer conformation in solution. In particular if diene enters the chain by Me-CH2 bond in nallyl complex it leads to formation of 1,4-units. If the same reaction occurs by CH-Me bond then 1,2-(3,4-)-units are formed. In the majority of cases diene acts as bidentat ligand coordinating on transition metal Me by both double bonds in cisoid conformation. The first connecting of diene leads to the formation of anti-complex which further can isomerize into syn-complex [ 175, 2081: CH=C% I \

,CH2

f&ns CH-i

CH: \

-

CH\

I

CHT

cH2anti-complex

(a)

,CH2

cy: :i

& & I l

PH - cH2 syn-complex

Anti-complex is responsible for cis- 1,4-polydienes formation. As a consequence cis- and trans-units content in chain is determined by ratio of chain propagation rate to the rate of anti-syn(cis-trans)-isomerization of end unit in AS. Authors of [ 1751 explain the formation of 1,2-(3,4-)units by coordinating polymerization mechanism when anti-syn-isomerization proceeds over o-state (a) in which o-bond is conjugated with double bond. Authors of [209-2131 suggest that polymerization proceeds on the crystal structure defects of salt MeX,,, i.e. on places where metal is on the surface. Under interaction of MeX,, with homogeneous organometallic component AIR3 in dependence on their chemical structure three main types of AS are formed: AS-1, AS-2, AS-3. They differ in geometry structure, configuration energy and relative structure. It was shown that relative parts of AS-1 - AS-3 even without consideration of kinetic factors correlate well with experimental yields of polydienes stereoisomers [2 10-2131. General regularity for heterogeneous Ziegler-Natta catalysts is unambiguously observed: AS-1 leads to the formation of trans- 1,4-polydienes; AS-2 forms high-molecular cis- 1,4polydienes; and AS-3 produces 1,2-(3,4-)-polydiens. Consequently stereoisomers ratio, in particular of cis-1,4 and 3,4-strustures under synthetic isoprene rubber synthesis in the presence of Ziegler-

28

Natta catalysts is determined by AS-2-to-AS-3 ratio, which in their turn is determined by catalytically active sediment structure. Under isoprene polymerization one of the important factors is MM and MMD control because strength and elastic properties of polymer in many respects is determined by MMD character. In particular too broad MMD means deteriorated rubber's properties. Prolonging of polymerization process initiated by lanthanide systems leads to MM increase and MMD change. However, when definite monomer conversion is reached (usually it is - 30%) chain propagation stops and further polymerization doesn't lead to essential change of molecular characteristics. The last fact is due to reactions of chain limitation in particular chain transfer on AlR3, monomer and other presenting in the system agents [174, 214, 2151. For Ziegler-Natta catalytic systems h4M reduction and h4MD broadening are characteristic when catalyst concentration increases. It is usually concerned with the fact that catalytic systems' components in particular organo-aluminum compounds take part in reactions of propagating chain limitation [215, 2161. However, if one increases the content of transition metal salt only at constant concentration of organo-aluminum component then molecular masses do not change [ 1741. As a rule polymer products synthesized on Ziegler-Natta catalysts have broad MMD that is concerned with the presence of several AS types differing in kinetic activity [217-2231. The method of determining of AS distribution function by possibility of propagating chain limitation based on total MMD data analysis and do not considering suggestions about its form was proposed for the first time under ethylene polymerization on titanium-magnesium catalysts [217]. It is accepted that each AS type produces polymer fractions Mh4D of which is obeyed the law Pi.exp(-PiM), Le. Flory's distribution [224], where p - the value inverse to M,,, characterizing possibility of chain

r-

1 1 termination and equal to -(rp- polymer chain propagation rate, ro - total rate of termination,

'6 m - monomer molecular mass). Resulting form of MMD is superposition of distributions formed at each polymerization AS. In [219-2231 method of determining of distribution of macromolecules propagation sites by kinetic activity was developed and intensified by the example of dienes polymerization in the presence of cis- and trans-regulating Ziegler-Natta catalytic systems. On the base of experimental MMD curves of polydienes with the help of Tikhonov's regularization method [225, 2261 distributions on kinetic activity of diene polymerization active sites for lanthanide [220, 22 11 and vanadium [222] catalytic systems were received. It was shown that even at low monomer conversions (about 1%) the obtained curves are polymodal. It is suggested that each maximum of distribution corresponds to at least one type of AS with definite possibility of propagating chain termination. This causes the existence of several AS types that take part in dienes polymerization initiated by Ziegler-Natta catalysts. Since Ti-AI, V-A1 and Nd-AI catalytic systems are micro-heterogeneous so the presence of various AS types differing in kinetic activity and stereospecificity [209-2131 is firstly adjusted with conception of absence of equipotentiality on catalyst solid surface [209-213, 227-2311. Defects of crystal lattice, heterogeneity, roughness of catalyst surface naturally lead to various kinetic activity and stereospecificity of catalyst. Moreover, heterogeneous catalytic system has definite catalyst' particles distribution on dispersion degree that determines dependence of their reaction ability under olefins and dienes polymerization [ 144, 229, 2321. It was shown in [230] that even porosity degree of heterogeneous catalyst surface influences on polymerization process kinetics. In [233] by the example of olefins polymerization authors explaine process rate change and MMD broadening by diffusion limitations under transportation of monomer to the active site across polymer film on catalyst surface. Thus, macromolecules stereoregularity is provided by chemical nature of used catalytic system and its formation conditions (components ratio, time ageing, temperature, catalytic system modification by electro-donors). At that by polymerization conditions varying one can influence on resulting polyisoprene MM and MMD. At the same time it is obvious that mechanical effect

29

(mixing) at the catalytic systems formation stage and under their mixing with reaction mixture influences on AS redistribution and polyisoprene microstructure because of heterogeneity of the majority of Ziegler-Natta catalysts [209-213, 227-2311. Changes in AS distribution on kicetic activity [218-2231 and consequently in resulting polyisoprene molecular characteristics are possible due to the same reason. It should be noted that under (co)polymerization of olefins and dienes initiated by ZieglerNatta catalysts process of bonding of the first monomer molecule to the catalytically active particle (initiation) differs from connection of the next monomer molecules (chain propagation) and proceeds so fast that there are no any data about constant of initiation rate [77, 174, 234-2361. Extremely high rates of AS formation and initiation of olefins and dienes (co)polymerization under process proceeding in stirred tank reactors [ 1761 propose diffusion limitations and determine possibility of appearance of heterogeneity of forming AS distribution as at micro-, so at macrolevels. Of course, this fact tells upon molecular characteristics of resulting polymer products [25, 78, 79, 236, 2371. Turbulent mixing intensification under polymerization sites formation and process initiation, i.e. diffusion limitations reduction is practically the only way of influence on the fast chemical processes character [ l ] . Moreover, it was shown in [238] that at polymer concentrations high enough (more than 3 mass %) intensity of mixing in reaction zone significantly effects on polymerization rate decreasing with the rise of viscosity. Authors propose to intensify catalyst mixing with monomer solution at least at the introduction into polymerizator where it can be reached practically without polymer to reach maximum polymerization rate. This suggestion was confirmed in industry conditions [ l , 25, 75-77]. When turbulence level increases in 5-10 times at reaction mixture formation moment under isoprene polymerization on Ti-A1 and Nd-A1 catalytic systems increase of cis-1,4-units content for 2-4%, reduction in 2-4 times of catalyst consumption, decrease of rate of crust formation on the inside surfaces of technological equipment in order were observed [75, 761. It is reasonable that investigation of turbulent mixing influence on process character and quality of resulting polymers by application of tubular turbulent apparatus at reaction mixture and microheterogeneous Ziegler-Natta catalytic systems formation stage under stereospecific isoprene polymerization seemed to be very expedient. 1.4.2. Pentadiene-1,3 cationic polymerization

Double-stage dehydration of isopentane is one of the major methods among industrial ways of isoprene production. Dehydration of isopentane and isopentene is carried out at high temperatures and is accompanied by formation of by-products. The most noticeable among them is pentadiene1,3 (piperylene) [239, 2401. Piperylene yield is 10-15% from diene monomers sum. One of the directions of piperylene application is cationic oligomerization for production of liquid rubber - the base of synthetic drying oil of SKOP brand. Use of piperylene hydrocarbon fraction as feed stock for synthetic drying oil production allows solving the problem of isoprene production by-product utilization and reducing consumption of plant oils in paintwork and other industry branches [206, 241, 2421. It is accepted that piperylene oligomerization in the presence of electrophilic catalysts (TiC14, A~CIYORZ, SnC14 and so on) proceeds according to traditional for cationic polymerization scheme. However, it was revealed in [240], that piperylene oligomers received in toluene medium contain 35 phenyl groups on 100 main piperylene units, i.e. toluene taken as a solvent take part in chain transfer reaction. This leads to molecular mass reduction under piperylene polymerization in benzene or toluene in comparison with paraffin hydrocarbons as solvents (900- 1200 against 12601430). Mechanism of chain transfer (TiC14.HzO catalyst) can be presented by following scheme

CH3

CH3

30

Ability of internal double bonds of forming macromolecule to further transformations initiated by cationic initiators is the characteristic particularity of cationic polymerization. This character leads to reduction of polydienes unsaturation due to the formation of brunched and cyclic structures even at low monomer conversions [ 1751. Total unsaturation of resulting polypiperylene usually is about 60-90% from theoretical [243]. One supposes [244] that these chain units consist of monocycle units of type 1-111:

I I1 I11 Structures I and I11 correspond to the most probable cycle units' formation. Moreover, decreasing of total oligopiperylene unsaturation may occur at the expense of macromolecules bonding [243] that should be taken into account when selecting catalytic systems for polymerization. Various catalytic systems both of ion [240] and ion-coordinated [174, 2451 types were proposed for piperylene polymerization. Polymer with low MM is formed under cationic polymerization due to the high rate of chain transfer reaction on monomer (it rises when catalysts acidity increases) and solvent (it reduces with the increase of polarity of solvent). Some halogenides of metals of 111-V groups were tested as catalysts of cationic polymerization. Tic14 and SnC14 turned to be the most useful. Application of SbCl5 and InC13 doesn't provide acceptable polymerization rate, and in the case of AlC13 insoluble polymer was formed [246]. Kinetic investigations presented in [247, 2481 showed that use of Tic14 as catalyst for industrial piperylene fraction polymerization (components content, mass %: trans-piperylene 60,8, cis-piperylene 37,2, etc.) allows to produce oligopiperylene with high yield (80-95%) with desired MM (80-2000) at comparable high process duration (nearly 1 hour). Polymerization process in toluene medium is characterized by the first orders by catalyst and monomer with constants of chain propagation, transfer on monomer and catalyst rates 1,5, 0,04 and 0,86 Vmolemin accordingly. Oligomers MMD has unimodal character, but with synthesis temperature rise products polydispersity increases. Results of study of piperylene oligomerization with the use of aluminum ethyl derivatives of various substitution degree by chloride-ion are presented in [244]. Independently on monomer isomer form unsaturation of oligomer chains is consisted of 1,4- and 1,2-trans-units as in the case of other electrophilic catalysts [206]. Relatively low-molecular products containing large amount of insoluble gel are formed in high yields under oligomerization of industrial piperylene fraction in solvent with use of C2H5AlC12 as a catalyst. The last fact causes the necessity of modification of with the aim of reduction of their acidic force. both applied catalyst RAlCl2 and In literature data there is an example of effective modification of organicaluminum catalysts by their combination with ethers [ 1751. Application of Al(i-C4Hg)C120Rz provides decreasing of pentadiene-1,3 oligomerization rate and elimination of gel formation reaction [240]. Complex formation of Al(i-C4H9)C12 with ethers leads to the decrease of Lewis' acidity of aluminum compound and value of acidic force decrease is determined by ethers electron-donor properties. In particular,

31

Lewis'

acidity

in

used

ethers

decreases

in

the

row:

A ~ R C ~ ~ ~ O ( C ~ H ~ ) ~ > A ~ R C ~ Z ~ ( C ~ H ~ ) ~ C H ~ > A L RConsequently, C~~~O(C~H one ~ ) ~can >A~RC~~ suppose that complexes activity under olefin or dienes oligomerization process will change in the same order [242]. It was shown in [249] on the base of quantum-chemical calculations and experimental data that acidic force of electrophilic catalysts is the measure of activity and stereospecificity of their action under cationic polymerization. Increase of acidic force of electrophilic polymerizations' catalysts promotes increase of their activity but decreases process selectivity. For example, in [250] piperylene oligomerization initiated by etherates of titanium chloride and aluminum is studied. AIC13.O(CsH5)2 gives high yield and molecular mass of product and received oligopiperylene however has low unsaturation. At present industrial piperylene oligomerization in the presence of electrophilic catalysts (TiC14+AlR3 [98], AlC13.O(C&)2 [250, 2511, etc.) for synthetic drying oil of SKOP brand production is realized in stirred tank reactor (12m3) or cascade (in 1,5m3).The last ones are equipped by mixers and external and internal coolers (coolant is reused water). Time of reaction mixture residence in apparatus reaches 3 hours. At the same time analysis of industrial production showed [252] that piperylene oligomerization in the presence of A1Cl3~0(C6H~)2 proceeded with high rates. Thus, Frydel-Kraft's catalysts on the base of aluminum or titanium halogenides with cocatalysts (HHal, H20, ROH, etc.) are widely used for piperylene oligomers synthesis. Moreover alkylaluminumchlorides possessing high activity at olefins oligomerization, having lower tendency to initiation of secondary reactions in comparison with aluminum halogenides and promoting gel formation in products are also perspective in respect to practical use. Furthermore active sites stabilization by ethers molecules getting as catalytic systems modifiers at the expense of salvation and oxonic ions formation also decreases their tendency to participating in secondary reactions, in particular macromolecules bonding. At the same time there are no data in literature on kinetic parameters of piperylene cationic oligomerization initiated by typical electrophilic catalysts. It is obvious that various electrophilic catalysts interesting for piperylene cationic oligomerization E2401 possess various activities. It determines principle possibility of improving of oligopiperylene industrial production as a base of synthetic drying oil production by selection of catalytic system with corresponding activity and selectivity. In this connection investigation of kinetics of piperylene cationic oligomerization in the presence of electrophilic catalysts, in particular TiC14, TiC14-Al(i-C4Hg)3, AlCzHsC12.O(C6Hs)2, A I C ~ H ~ C AlC13.0(C6H5)2, ~Z, Al(CzH5)2Cl and kinetic calculation of reaction zone (reactor) geometry are seemed to be expedient.

1.4.3. Ethylene and propylene copolymerization Ethylene-propylene rubbers are the important class of elastomers and are widely used in many branches of industry due to their unique properties, in particular good stability to environmental effects, good physical-chemical and elastic characteristics, chemical and ozone-stability, radiation, oils, acids and alkalis stability, etc. [78, 79, 176, 207, 2531. Triple ethylene-propylene rubbers of SKEP(T) brands contain up to several percents of nonconjugated dienes predominantly dicyclopentadiene (DCPD) and ethylidennorbornene (ENB) as third comonomers [ 1771. At the first time ethylene and propylene copolymers were put into industrial production in 1959 in Italy due to Ziegler-Natta catalytic systems discovery [176, 1771. At present various SKEP(T) types are produced in all leading in industry countries particularly in USA, Italy, Germany, Holland, Great Britain, Canada, Japan, France, etc. Nevertheless, there is a lack of ethylene-propylene rubbers at world markets at present and this fact determines advisability of their production expansion. In Russia the first and the only large production of SKEP(T) in high-boliling hydrocarbons medium (hexane, heptane, etc.) was placed in operation on zavod SK OAO

32

"Nizhnekamskneftekhim" in 1995-1996 [78, 791. Polymerization in hydrocarbons solvents is carried out at 20-40°C and pressure 0,2-1,4 MPa. Polymerization duration is 03-13 hours. Disadvantages of existent method of SKEP(T) production are: problem of effective heat removal; limitations of polymer content in solution (from 10 to 12%) due to significant rise of viscosity; high power inputs for regeneration of high-boiling solvent; irregular saturation of liquid products by monomers and hydrogen prior to gas-liquid mixture introduction into the reactor; reception of irregular in structure copolymer both in one reactor and parallel operating polumerizates. Among known Ziegler-Natta catalytic systems catalysts on the base of V and Ti compounds combination with chloroaluminumalkyles are effective for ethylene and propylene copolymerization [176, 1771. It is particularly convenient to use systems on the base of vanadium compounds (tetrachloride, trichloroxide, triacetylacetonate) and diisobutylaluminum chloride. Under ethylene-propylene rubbers production in the presence of Ziegler-Natta catalytic systems as well as under isoprene stereospecific polymerization (see 1.4.1) the rates of active sites formation and initiation are essentially higher than chain propagation rate [81, 174, 234-2361, This fact determines heterogeneities of polymerization AS distribution over reactor's volume. One revealed that in this case it was possible to use the prereactor of divergent-convergent design with positive effect at catalytic complex formation stage and its introduction into reactor [77, 254, 2551. Prereactor was commercially tested and applied on A 0 "Kauchuk" (Sterlitamak city) from 1996 for isoprene stereospecific polymerization with rubbers of two brands SKI-3 and SKI-5 as results [75, 761. Characteristic feature of ethylene-propylene rubbers synthesis is that the process proceeds in two-phase system and transportation of two monomers (ethylene and propylene) to the liquid before chemical interaction. Ethylene-propylene rubber is irregular in molecular-mass structure as any other polymer synthesized by this way [lo, 2561. First of all this irregularity is explained by difference in monomers concentrations in solution and directly in the reaction zone. In [236, 237, 257, 2581 authors showed that with increase of viscosity with molecular mechanism of transfer rise of diffusion resistances was expressed in the increase of liquid film thickness. The last fact leads to the appearance of bimodal MMD. Moreover, the presence of "gel-fraction" in polymer is caused by formation of macromolecules with high content and block degree of ethylene units [78, 791. For example, if ethylene units content in normal polymer is about 57-60%, then in hardly dissolved fractions (in gel) - 79-83%. It is obvious that reception of heterogeneous in structure copolymer including that with high "gel" content is determined by heterogeneous distribution of liquid products, monomers and hydrogen before mixture introduction into polymerizator, i.e. by low level of turbulent mixing of gas-liquid flows in stirred tank reactors as in the case of MMD broadening. Significant increase of rate of gases dissolution in liquid while using tubular turbulent apparatus [ 110, 2591 determines expediency of their application in production of ethylene-propylene rubbers at the stage of saturation of solvent by monomers. At present tubular turbulent apparatus system of divergent-convergent design is used in ethylene-propylene rubbers production at the stages of catalytic system decomposition, removing of catalyst from polymer and introduction of stabilizer-antioxidant into rubber solution [78, 791. Thus, the main problem of SKEP(T) production requiring solution is the rise of turbulent mixing at stages of liquid and gas flows mixing prior to their introduction into parallel operating reactors-polymerizators, and at the stage of formation of sites of macromolecules propagation that determines possibility of production of qualitative and homogeneous in structure copolymer. 1.4.4. Chlorination of butyl rubber

Halobutyl rubber is one of the butyl rubber (BR) modifications (chlorinated (CBR) and bromated (BBR) rubbers) the main advantage of which is its ability to vulcanize with any types of

33

rubbers [207]. Due to the last fact demand for halobutyl rubber is constantly increased whereas demand for butyl rubber is reduced. Halobutyl rubber is an excellent raw material for tyre industry: tubeless tyre, high-temperature and conveyor belts, etc. [143, 2601. Halobutyl rubbers part in total BR production abroad is about 40%. In Russia and other countries of near-abroad there is no industrial production of halobutyl rubbers including chlorinated butyl rubber [55, 100, 2611. Macromolecules of CBR consist of isobutylene orders fragments (blocks) that are statistically divided by chlorine-containing isoprene units of various structure [27, 176, 2621. Under BR chlorination total unsaturation of macromolecules and especially content of internal C=C bonds are decreased that is easily defined by split ozonolysis [263]. Reaction of BR chlorination is occurred according to scheme [27,263,264]: -iw2-qq=cH-cH2-

a2

CBR molecules predominantly contain exomethylene groups [structure (2)] - up to 90%. Content of endomethylene groups [structure (l)] is equal to 9%. Saturated groups [structure (4)] can also present in macromolecules structure - up to 1%, that are predominantly formed at low temperatures [27, 176, 2651. Relative structures content 2 : 1 determines CBR performance attributes in particular thermal load stability and stability under processing [ 1431. Up to introduction of 1 chlorine atom on 1 isoprene unit under BR chlorination increase of chlorine content in BR is observed (linear dependence). 50 mole % from introduced C12 enters into BR. Process of BR chlorination should be accurately controlled. Molecular mass of CBR is reduced not more than by 5 1 0 % under introduction of up to 1 chlorine atom on C=C bond. Under more deep chlorination significant reduction of C=C bonds content in macromolecules and proceeding of deep CBR destruction down to liquid products formation are observed. It is possible to make less severe demand for temperature under BR chlorination because in 10-55OC diapason temperature change in reaction zone doesn't essentially influence on process and quality of resulting CBR. However CBR intensive destruction is carried out if temperature is higher than 55°C. Among famous methods of BR chlorination by molecular chlorine or chlorine-containing organic compounds in solution, water dispersion and liquid melt [27, 143, 176,262, 2661 chlorination of BR by molecular chlorine in solution is the most manufacturable method. Process consists of several power-consuming stages [27, 176, 1771: 1) preparation of 10-15% solution of BR in aliphatic (hexane, benzene fractions) or chlorinated (methylchloride, methane tetrachloride) hydrocarbons; 2) chlorination of BR by molecular chlorine in solution; 3) neutralization of resulted CBR solution; 4) removing of salts, etc. by water out of CBR solution; 5) introduction into CBR of stabilizer-antioxidant (0,38 from rubber mass) and antiagglomerator calcium and zinc stearates (0,8-1,2%); 6) CBR water degassing; 7) polymer isolation and drying; 8) removing of solvent, its azeotrope dewatering and rectification of returned solvent. Tank apparatus equipped by mixing and in some cases heat-exchanging devices, intensive mixers and bulk washing columns are used practically at all process stages

34

Necessity of use of high viscosity 10-15% rubbers solutions (CBR and BR) in organic solvent (p is about 0,3-0,6 Pa.sec) is specific feature of this process. Effectiveness of each stage (in particular stages 2-5) depends on mass-exchange effectiveness that in the case of high viscosity solutions is a serious problem. Application of gas substances - mixture of molecular chlorine and nitrogen (usually in volume ratio 1 : 6 - 1 : 5) is the particularity of CBR synthesis. This fact determines the situation when gas mixture volume is practically in 10 times and more higher than volume of high viscosity BR solution in organic solvent. So, special attention is required under creation of optimal conditions for BR chlorination in liquid-gas system, firstly at realization in reaction zone of finebubble (foam) regime under "gas-viscous liquid" flows mixing. So, rotary pump as additional apparatus-chlorinator into which BR solution and chlorine with nitrogen mixture are introduced is widely used for the intensification of BR chlorination process in stirred tank reactors. Recycle is stipulated for providing of necessary residence time of reacting mixture [267]. Thus, under industrial CBR production one should solve several complex problems: - creation of optimal conditions for effective mass-exchange in "viscous liquid-gas'' system, in particular formation in reaction zone of fine-bubble (foam) regime: - providing of such conditions for BR chlorination in solution when characteristic mixing time zniXwill be lower (or at less comparable) than characteristic time of chemical reaction Tch; - providing of optimal conditions for homogeneous emulsions formation under combination of BR solution with water with proceeding of chemical reaction (neutralization HCI) or physical mass-exchange processes (removing of salts, etc. out of CBR solution, introduction of stabilizers into CBR).

35

CHAPTER 2. EXPERIMENTAL PART 2.1. Experimental work 2.1.1. Oligomerization of piperylene

Experiment was carried out at intensive mixing in adiabatic conditions at initial temperature 298K in glass flask of volume equal 100 cm3 equipped by thermometer for measuring of temperature rise under vacuum and moistureless conditions. Catalyst was introduced into flask by syringe and this moment was considered as polymerization beginning. Reaction was stopped by methanol. Then polymer was isolated by methanol containing 1% of stabilizer - ionol, thoroughly purified from catalyst leavings and dried at 313K in vacuum until its weight became constant. Relative experiment's error was 2-3%. In the course of piperylene oligomerization reaction mixture temperature was measured. Kinetic curves of monomer conversion dependence on time were obtained out of dependence of reaction mixture temperature adiabatic rise AT on polymer yield An: AT = qAn/C,p (2.1) For calculation of chain propagation rate constant initial parts of kinetic curves were used, when temperature of reaction mixture did not markedly differ from initial. So, received values of piperylene oligomerization rates constants in first approximation may be referred to temperature interval 298-3 10 K. Relative methods error was 15-20%.

2.1 -2. Isoprene stereospecific polymerization The scheme of experimental assembly for isoprene polymerization in toluene is presented in Figure 2.1, Initial reagents were purified using methods conventional for stereospecific polymerization in the presence of Ziegler-Natta catalysts. At equal conditions of polymerization [Ti : AI : piperylene = 1 : 1,25 : 2 mole (piperylene was introduced as modifier into catalytic complex solution at the moment of its preliminary ageing), ageing of catalytic system for 30 min at 273K, polymerization temperature 298K, toluene as a solvent, CTi= 6 mmole/l, CM= 1,5 mole/l] way of process carrying out was changed. Namely: Variant 1. Preliminary aged catalytic complex (two-component TiCl4-Al(i-C4H9)3 and ternary TiCldAl(i-C4H9)3-piperylene) and solvent from reservoirs 1 and 2 (Fig. 2.1) accordingly were mixed in tubular turbulent apparatus 3 at linear rate of flows in broad part (divergent) not less than 0,5 d s e c and reagents residence time in mixing zone nearly 2-3 sec (hydrodynamic effect on catalytic system in turbulent regime). From apparatus 3 catalyst solution entered into vessel 4 where isoprene was added. Then polymerization was carried out at slow mixing. Variant 2. Solutions of Tic14 and Al(i-C4H9)3 from reservoirs 1 and 2 accordingly through apparatus 3 entered into vessel 4 where resulting catalytic complex was aged for 30 min at 273K (catalytic system preparation in turbulent regime). Variant 3. Solutions of preliminary aged ternary catalytic complex and monomer from reservoirs 1 and 2 accordingly were mixed in apparatus 3 with following introduction of reaction mixture into apparatus 4 (reaction mixture formation in turbulent regime). Traditional conditions for isoprene polymerization were created in tank apparatus (vessel) 4 at low mixing by magnetic stirrer (100 rpm) and tubular turbulent apparatus in this case functioned as prereactor (Fig. 2.1). For each of noted variants blank experiment modelling traditional process scheme was simultaneously carried out (initial reagents without preliminary mixing in prereactor 3 were introduced directly into reaction volume 4). Conversion curves of isoprene polymerization were received by gravimetric method.

36

Figure 2.1. Experimental assembly for isoprene polymerization. 1, 2 - reservoirs for reagents; 3 tubular turbulent apparatus; 4 - vessel of 500 cm3 volume; 5 - three-way cock; 6 - magnetic stirrer.

2.1.3. Chlorination of butyl rubber

Experiment was carried out on assembly of experimental-industrial shop of research center of OAO "Nizhnekamskneftekhim". Butyl rubber has been preliminary dissolved in nephras in reservoir with stirrer for 5 hours up to solution concentration 13%. Then solution with the help of the pump was inputted under pressure 0,08 MPa at 303K to the connecting nipple of tubular turbulent apparatus of divergent-convergent design with following sizes: divergent diameter dd = SOmm, convergent diameter d, = 25mm, section length L, = 75mm, L = 2m. Connecting nipple was equipped by nozzles allowing receiving of fine disperse emulsion with constantly renewed surface. Chlorine and nitrogen gas mixture was preliminary prepared in ratio 1 : 5 and under pressure 0,1 MPa, then it was inputted to the corresponding reactor jet; at that ratio of chlorine to double bond was equal to 1,2. Gas-liquid mixture outgoing from reactor was inputted to separation after which rubber solution was purified by water, and then by 10% solution of NaOH. Chlorobutyl rubber solution purified from salts was degassed by direct steam. Estimation of chlorine content values dispersion in chlorobutyl rubber was held on three samples taken every 30 minutes.

2.1.4. Ethylene and propylene copolymerization Experiment was carried out on assembly of experimental-industrial shop of research center of OAO "Nizhnekamskneftekhim". Gas-liquid mixture cooled down to 263K was inputted into two reactors-polymerizators of 16,6m3 volume both with speed of rotation equal to 35rpm. Gas-liquid mixture had following components ratio, mass parts: Liquid propylene 0,075 Ethylene 0,05 Hydrogen 0,025 Recycling gas (ethylene, propylene, hydrogen) 0,35 Nephras 0,s The third monomer (DCPD, ENB) was introduced in amount of 75kg per ton of copolymer.

31

Components of gas-liquid mixture were inputted by doses into two sectional consecutively connected tubular turbulent apparatus of divergent-convergent design with dd = 150mm, d, = 75mm, L, = 450mm, y = 45". Convergent of the last fifth section of the second tubular apparatus was connected with nozzle for gas-liquid mixture taking off to the parallel operating reactors-polymerizators. Sectional tubular turbulent apparatus of divergent-convergent design with dd = 80mm, d, = 40mm, L, = 200mm, y = 45" were placed ahead of gas-liquid mixture inputting nipple. Solutions of catalytic complex components in nephras [VOC13 and Al(CzH5)2Cl] were inputted directly into each polymerizator counting on 2% from gas-liquid mixture volume. Pressure inside of reactor-polymerizator was 0,45MPa, temperature 308-3 18K, polymerization time 03- 1Shour. Copolymer was purified, stabilized and briquetted after process completion. 2.1.5. Formation of macro-structures of reaction and mixing fronts in turbulent flows

Mixing of liquid flows was carried out in glass tubular turbulent apparatus of cylindrical construction of two types that differed in ratio of internal (axial) sucker diameter dl to apparatus tubular canal diameter d3. Apparatus 1 - di = 10,2mm, dz = 1 3 m , dd = 23mm (di / dd = 0,44). Apparatus 2 - dl = 3mm, dz = 13mm, dd = 23mm (dl / dd = 0,13). Conditions of mixing fronts' macrostructures formation were studied on model assembly (Fig. 2.2) by "coloured flow method". Through the axial sucker 1 with diameter dl coloured flow of water solution of methyl green or potassium permanganat was entered and through side sucker 2 with diameter dz non-coloured flow (water) was inputted. Reagents solutions interacting with various reactions rates constants and resulting in formation of coloured products were entered into tubular apparatus for creation of reaction fronts' macrostructures. Chemical reactions of the second order were used: HC1+ NaOH + NaCl + HzO (k = 10'' l/mole.sec)

2

Fe(SCN)3 + 3KC1 3KSCN + FeCl3 (k = lo4 l/mole.sec) 3KSCN + FeC13 + Fe(SCN)3 + 3KC1 (k = lo2 l/mole.sec) (k = 10' l/mole.sec) &[Fe(CN)6] + FeC13 + me[Fe(CN)6] + 3KC1 Calculated amount of glycerin was added into solutions for obtaining of necessary viscosity. Soluble in water salts, in particular sodium chloride, were added into solutions for obtaining of necessary density. Rotameter reading controlled the speed of liquid flows inputting. Formation of plan front and torch in dependence on linear speeds of axial VI and radial VZ reagents flows introduction was fixed by sight and rapid electronic filming method (video camera Hitachi VM 2980E, focus 1211; digital photo camera SONY DSN - 75 (exposure 1/1000sec, photosensitivity 400)).

38

5

4

d

2

I

Figure 2.2. Model assembly for studying of liquid flows mixing process. 1, 2 - suckers of flows inputting; 3 - tubular turbulent apparatus; 4, 5 - water line; 6, 7 - reservoirs; 8 - pump; 9, 10 mixing block; 11, 12 - rotameters. 2. I.6. Suspensions preparation The following fast chemical reaction was used as a model: HzS04 + BaClz + BaS04&+ 2HC1 Three methods of suspension synthesis were used: Variant 1 (Fig. 2.3a). Barium chloride interacts with sulfuric acid placed in stirred tank apparatus (three-necked flask of 500cm3 volume and filled for 70%) in the presence of rotating mixer 2. The last one operates according with rotary pump principle and creates intensive turbulent flows in axial direction of reaction; Variant 2 (Fig. 2.3a). This method is analogous to variant 1. The only difference is that in reaction mixture volume baffle wall 1 is placed that along with increase of power spent on mixing intensifies turbulent flows mixing; Variant 3 (Fig. 2.3b). Interaction of initial reagents (water solutions of HzS04 and BaCh) was carried out directly in six-section tubular turbulent apparatus of divergent-convergent design 5 (divergent diameter is 24mrn, convergent diameter is lOmm, section length is 48mrn) at linear speeds of reagents flows movement V 4 , 5 d s e c . With the aim of increase of ratio nucleation rate / crystal growth rate and consequently with the aim of obtaining of maximum possible disperse systems concentrated (0,52 mole/l) and diluted (0,03 mole/l) water solutions of BaC12 and HzS04 accordingly were used.

39

urn

Figure 2.3. Experimental assembly for suspension synthesis. 1 - baffle wall; 2 - mixer operating by rotary pump principle; 3 - mixer's electric drive; 4 - three-way cock; 5 - tubular turbulent apparatus; 6, 7 - reservoirs for HzS04 and BaCl2 solutions, accordingly. 2.1.7. Emulsions preparation Experimental assembly for studying of regularities of dispersion in turbulent flows is presented in Figure 2.4. Glass tubular turbulent apparatus differing in canal geometry and method of flows inputting were used (Table 2.1). The water with consumption w = 40-200 cm3/sec was used as dispersing medium, hexane with consumption 1-13 cm3/sec was a dispersing phase. Frequency curves of distribution of disperse phase drops sizes were obtained by photography method at intensive transmitted illumination with the use of digital photo camera SONY DSN-75 (exposure 1/1000 sec, photosensitivity 400).

Figure 2.4. The scheme of experimental assembly of emulsification in turbulent regime. 1 - reservoir for disperse phase (hexane p = 0,66 g/cm3, = 0,3 mPa.sec), 2 - peristaltic pump, 3 - disperse phase line (water), 4 - rotameter, 5 - tubular turbulent apparatus of divergent-convergent (cylindrical) design (Table 2. l), 6 - digital photo camera, 7 - light source, 8 - collector.

Rapid photography reduces to minimum drops image softening effect due to high movements rates. Zinc-coated calibrating threads of known size (diameter - 0,2, 0,5 and lmm) were introduced into apparatus body for taking into account of image fault due to bend shape of apparatus [18].

40

Central part of apparatus longitudinal section with size equal to 0,75 x dd was used for calculations where discrepancy between calculated by obtained image and real calibrating threads sizes was less than 1%. This method also allowed excluding of possibility of image fault as a result of refraction coefficient of apparatus glass body and consideration of accuracy error. We obtained image of all particles in flow's longitudinal section because we have worked with disperse phase's inclusion volume fractions small enough (disperse phase consumption was 1-1,5cm3/sec, and consumption of disperse medium was w - 40-200 cm3/sec). Table 2.1. Characteristic of six-section tubular turbulent amaratus (& = 24mm)

9

Cylindrical construction with length L = 0,6 m

Frame-accurate review of moving disperse systems computer image (24 pictures per second) with tenfold magnification under confrontation of disperse phase particles linear sizes (all particles in selected flows movement rates diapason had spherical shape) with linear sizes of calibrating threads allows calculating of true drops' size (rather the area of spherical drop projection). Calculation of drops was held in following intervals (mm): 0-0,4; 0,4-0,6; 0,6-0,8; 0,8-1,O; 1,O-13; 13-2,O; higher than 2. Estimation of distribution dispersion s2 was made by experimental data:

A

-

where N - number of measurements; di - value of drops' diameter in selected interval; d - average value of drops' diameter. Standard error A at confidence probability a = 0,95 was calculated according with A = f2,4-.

S

6

Average measurement error was &7%.

Due to fine-disperse emulsions reception at high dispersion medium movement's rates w and under apparatus No.8 use (Table 2.1) microscopic method was used in this work in spite of photograph analysis of disperse inclusions. Microscopic method was realized with the help of LOMO-MIKMED-1 microscope with screw ocular micrometer EFM-16X-E (twenty-fold increase). Microscope method is described in [108]. In the work lower limit of disperse phase diameters for photographic method was about 0,2 mm, and the high level under microscopic analysis was 0,5 mm. Results of two methods of obtained emulsions analysis correlated with coefficient not less than 0,95 (Fig. 2.5), that allowed investigation of received emulsions in interval of disperse inclusions' sizes broad enough (diameter from 0,Ol to 3 mm).

2.1.8. Studying of hydro-dynamic structure of reaction medium motion in tubular turbulent apparatus Scheme of experimental assembly for hydrodynamic structure investigations is presented in Figure 2.6. Tubular turbulent apparatus of two types were used: cylindrical and of divergentconvergent design with parameters presenting in Table 2.2 (Fig. 2.7). Impact introduction of inert

41

indicator (methyl green) (2 cm3) into feeding flow with photometric measurement of its concentration Ci at apparatus output was carried out for reception of response curves (Fig. 2.7). Concentration measurement was carried out from the moment of indicator introduction till limit of its detection. Curve of indicator solution light-absorbing obtained on spectrophotometer "KPK2Mp" has its maximum at wave length h = 540 nm. Calibration curve (A= 540 nm) was made by series of standard solutions with indicator concentration 10-6+5.10-4mole/l (Fig. 2.8). The equation of calibration curve was obtained: D = 0,1139Ci + 0,0058 (correlation coefficient R=0,99), it was used for calculation on indicator concentration in outputting flow. ds, ~

l

l

1

o'6

0.4

-

0.2

-

0.J

i

40

60

80

100

120

140 w* cm3bec

Figure 2.5. Plots of surface-volumetric diameter d32 versus volume flow of disperse medium we. Apparatus No.8 (Table 2.1). Photography ( A )and microscopy ( 0 ) analysis methods of dispersion inclusions' sizes.

1

I

4

-

3

-M2 Figure 2.6. Experimental assembly scheme for investigation of hydrodynamic flows structure in tubular turbulent apparatus. 1 - tubular turbulent apparatus, 2 - water line, 3, 4 - rotameters of central (w1) and side (w2) flows, 5 - block of indicator introduction, 6 - spectrophotometer.

42

Idd

W +

4

x

Id n W

----+

a

I dc

c

II

Figure 2.7. Tubular turbulent apparatus of cylindrical (a, c) and divergent-convergent (b) designs with coaxial (a, b) and radial (c) methods of reagents introduction. Ind - introduction of indicator. Reactors working regimes: P - plan front; T - torch; Z - drift regime. D

0

2

4

6

0 C,XIO’,

10 moleil

Figure 2.8. Plots of solution’s optical density D versus indicator concentration Ci.

43

Table 2.2.

2.1.9. Investigation of convection heat-transfer Glass tubular turbulent apparatus made by analogy with heat exchangers of "annular tube" type with internal canal of cylindrical (d = 16mm) and divergent-convergent (Fig. 2.9) (divergent diameter dd = 16mm, convergent diameter d, = 8mm, section length Ls= 32mm) designs with length equal to l m were used for modelling of heat-transmission under fast exothermal chemical process with external cooling. Water was used as heat-transfer agent with various methods of flows movement (forward or back flow). Mercurial thermometers were used for water temperature measuring at input and output of internal and circular canals (Fig. 2.9). Distribution of flow's particles by residence times zr in heat-exchange zone with the help of curves of response to impulse inert indicator introduction was used for studying of hydrodynamic structure of reaction mixture and coolant movement in tubular turbulent apparatus' canals (Fig. 2.9) (see 2.2.8).

Figure 2.9. The scheme of experimental assembly for convection heat-transfer investigations. 1 thermometer, 2 - rotameter, 3 - tubular turbulent apparatus of cylindrical (divergent-convergent) design with jacket, 4 - spectrophotometer, 5 - manometer. Ind - indicator input, I - internal flow, I1 - circular flow.

2.2. Analysis methods

2.2.1. Determination of polymers microstructure Microstructures of piperylene oligomers' unsaturated part and polyisoprene were determined by IR-spectroscopy on "Specord M80". Absorption bands at 730, 750, 910 and 967 cm-' corresponding to cis-l,2-, cis-l,4-, 3,4- and trans-( 1,2 + 1,4)-units of oligopiperylene accordingly were used as analytical. Polyisoprene microstructure was determined by absorption bands 840 and 889 cm-'

44

corresponding to mixture of 1,4-cis + 1,4-truns and 3,4-units, accordingly [268]. Optical densities of analytical bands were determined by "basic" lines methods.

2.2.2. Determination of polymers molecular-masses characteristics Oligopiperylene solutions' intrinsic viscosity was determined in toluene at 303K in Ubbelode's viscosimeter with hanging level by extrapolation of qspI C and ln(qrelI C) values to zero concentration [269]. Molecular mass and molecular-mass distribution of piperylene oligomers were determined by gel-chromatography on chromatograph KhZh-1302 with the use of styrogel columns. Chloroform was the eluent. Polyisoprene molecular mass and molecular-mass distribution were determined by gelchromatography on chromatograph "Waters" with styrogel columns with pores sizes (A) 5.1O3-1,5+1O4.Toluene was the solvent, temperature 303+0,5K, elution rate - 1 cm3/min. Molecular-mass distribution of SKEPT samples was determined by spectra of relaxation times of polymers solutions pressure (SRTP). Method is based on the use of experimental data of pressure fall in capillary viscosimeter's cylinder under non-stationary polymer melt outflow through capillary after piston stop (automatic capillary viscosimeter MPT "Monsanto" with capillary size n = 1Smm at 398K and initial shear rate 3,6 sec-') [270,271] .

2.2.3. Determination of chlorine content in rubber Chlorobutyl rubber sample was burned in oxygen atmosphere in the presence of platinum catalyst with the following absorption of burning products by alkali. The amount of formed halogenide-ions was determined by mercurrometric titration (titrant - Hgz(N03)2, indicator Fe(SCN)3) [272].

2.2.4. Determination of double bonds content in polymers The content of double bonds in polymers was determined by split ozonolysis at 273K on analyzer of double bonds ADS-4 ICPh RAS.

2.2.5. Determination of density and viscosity of liquidjlows Determination of density of liquid flows inputting into tubular turbulent apparatus was carried out by bottle method according with all-Union State Standard 3900-85. Viscosity determination of flows was carried out on viscosimeter VZ-249 according with all-Union State Standard 9070-76.

2.2.6. Analysis of suspension dispersed structure Dispersing analysis of resulted Bas04 and Ziegler-Natta catalytic systems suspensions (see 2.2.2 and 2.2.6) was carried out by sedimentation method in gravitational field with the use of torsion balance according with standard technique [273]. Dilute suspensions (-O,35 % mass) were prepared for decreasing of possibility of particles interaction in the course of precipitation. Graphical differentiation of sediment accumulation curve was used for sedimentation analysis data handling. For this sedimentation curves of formed sediment m dependence on time z were approximated by sigmoid three-parameter function:

45

a

m(d =

(2.3)

2-X

I+ expt 01

b

Mixing rate N, rpm 300 1500 2500

a

b

xo Stirred tank reactor with baffle walls (Variant 2) 295,1+2,1 6,6+0,2 32,4+0,2 289,8+3,2 14,2+0,6 72,8f0,8 293,3+1,8 19,lf0,6 88,9+0,8 138,7+1,4

30,7+1,1

303,2+2,6

Correlation coefficient R 0,99 0,99 0.99 0,99

In this case according with Svedberg-Oden's equation [273] mass mi of fractions of particles with radius r > ri participated at time moment Ti was calculated by the following expression: 2

mi=

expt -)a

B -2

1+ expt-0)

-x

b

--2

i

-x

b

-x b(l+e x p C u ) ) 2 b -2.

(2.4)

Since differential distribution of particles of disperse phase F(re) by equivalent radius re (radius of spherical particle precipitating with the same rate) is typical function of variable distribution, so for calculation of most probable (averaging by mass) radius rmpof suspension particles the following expression was used: W

46

CHAPTER 3. DEVELOPMENT OF SCIENTIFICALLY-GROUNDED APPROACH TO THE SELECTION OF OPTIMAL REACTION ZONE GEOMETRY UNDER FAST PROCESSES Creation of high turbulence in reagents mixing zone without preliminary flows turbulization is necessary for the effective fast chemical reactions and mass-exchange physical processes proceeding in liquid phase. As a consequence, studying of fundamental regularities and revealing of quantitative dependences allowing creation of intensive turbulent mixing in reaction zone are very important stages of novel technologies development. One should consider processes occurring in homogeneous mediums and under phases interface presence conditions for the solution of this important scientific and applied problem.

3.l.One-phase systems 3.1.1. Mathematical modelling

Theoretical description of reagents turbulent mixing in tubular canals is based on the following main model assumptions: - the medium is Newtonian and incompressible; - the flow is axisymmetric and nonswirling; -turbulent flow can be described by standard model [161] with following parameters: kinetic energy density of turbulence K, dissipation rate of turbulence E ; - coefficient of turbulent diffusion is equal to kinematic coefficient of turbulent viscosity D, = V, = b / p. Then they choose numerical solution of continuum turbulent flow equations with effective viscosity coefficient p = b + pM (pM- coefficient of molecular viscosity) by the use of IC-E turbulence model with method of finite elements on irregular computational grid [274]. These equations in cylindrical co-ordinates are as following: Equation of continuity - - + - = o , av r ar az Nav’ye-Stocks’ equation (equation of impulse transfer):

lam)

where stress tensor components are equal to:

Equations of transfer of kinetic energy density of turbulence and its dissipation:

r

ar

47

+p.C GE/K-C$E~/K, 0 1

C K2 cLT=P-9

CL

E

(3.9) where p - pressure; r, z - radial and longitudinal coordinates, accordingly; u, v - cross and longitudinal rate's components. Standard parameters of IC-& turbulence model were used [274]: C1 = 1,44, CZ= 1,92, C , = 0,09, UK = 1,0, UE = 1,3. Equations (3.1)-(3.9) are true for the main liquid volume but cannot be applied for the region directly adjoining to reaction zone's wall. For the last case one should use wall's law according with which the flow speed profile close to solid wall obeys the logarithmic law: (3.10) where 6 - the distance from the wall. In the course of iterative process at given rate of reaction mixture coaxial movement Vtanone calculates the tangential stress r by (3.10) and then the effective viscosity: CL = h N t a n , (3.1 1) by this the turbulence kinetic energy K is calculated from IC-& turbulence model, and dissipation of kinetic energy density of turbulence E as follows: (3.12) For formulas (3.10)-(3.12) constants values are accepted k = 0,4, E = 9,O and this is the standard boundary condition for all solid surfaces under turbulent flows [161]. Flow of liquid with viscosity coefficient pM= 1 mPa.sec and density p = 1000 kg/m3 was considered. Boundary conditions are the symmetry conditions along z axis and conditions of liquid adhesion to solid surfaces of reaction volume. They set pressure at apparatus output (on CD line) and linear flow speed V = 5 d s e c at input (on AB line) in the line of symmetry axis (Fig. 3.1). The lengths of input and output of reaction zone significantly exceed zone's diameter (L >> dd) that allows exception of influence of input and output turbulence parameters on reagents mixing characteristics. The last ones are the subjects of inquiry. They compared obtained theoretical results with available experimental data for reaction zone of cylindrical type to confirm adequateness of carried out calculations (Fig. 3.la) (Re = 2.105) [275]. By this the received results are agree with calculations made in [126]. In particular, values of length Lfir of circulation zone appearing in peripheral regions of reaction zone directly after reagents introduction into extending canal are presented in Table 3.1. There is a good coincidence (error is not higher than 15%) of calculated and experimental values of circulation zone length.

48

a

C . . . . . . . . . . . . . . . . . . . . L A

C

b

A

I

PEG I1

C I11

N

Figure 3.1. The schemes of cylindrical (a) and divergent-convergent (b) apparatus. y - the angle of divergent opening; Rd=dd/2 - radius of wide part (of divergent); R c = de12 - radius of narrow part (convergent); 1-8 - apparatus parts; I-IV - divergent-convergent sections.

Lcir 1 dd

The angle of divergent opening y, degrees

Experimental data [275]

30 90

41 46

Calculation Our results [I261 3,4 4,7

3s 4,7

t

Figure 3.2. The profile of axial rate in section z = dd of cylindrical apparatus at y = 30' (a) and y = 90' (b): - experimental data [275]; dotted line - calculation data [126]; firm line calculation data (our results). Thus, chosen mathematical model allows calculating of characteristics of turbulent mixing in tubular apparatus with various geometry of canal. Adequacy of obtained results (see Fig. 3.2 and 3.3) is well confirmed by correlation of experimental and calculating data.

3.1.2. Turbulent mixing in tubular apparatus Numerical calculations on the base of (3.1)-(3.10) equations allowed studying of regularities of turbulent diffusion coefficient Dt variation along the volume of tubular turbulent apparatus. Identification of characteristics of turbulent mixing in tubular turbulent apparatus with local hydrodynamic resistances of divergent-convergent design with optimization of their geometric parameters is of the main importance. Calculations showed that when divergent angle of opening y increases from 5 to 3OoC, i.e. at transition from cylindrical apparatus (Fig. 3. la) to divergent-convergent design (Fig. 3. lb) the coefficient of turbulent diffusion rises in 3 times. But with further increase of angle the coefficient practically doesn't change (Fig. 3.4). The constancy of turbulent mixing level in apparatus peripheral and central regions is noticeable and it determines the equality of reagents turbulent mixing characteristics along the whole apparatus volume in wide interval of values of divergent angles of opening (Fig. 3.4). Firstly, this fact due to significantly lower convection rate of liquid flows in peripheral region (R > dJ2) than in central part (R < dJ2) of tubular turbulent apparatus of divergent-convergent design (Fig. 3.5). As a consequence the averaging of Dt along tubular apparatus volume at the expense of transfer processes predominantly occurs at turbulent exchange. EN2 0.04

0.02

0

0.4

0.8

rL&I

Figure 3.3. Profile of kinetic energy density of turbulence in the section z = & of cylindrical apparatus at y = 90': - experimental data [275]; dotted line - calculation data [ 1261; firm line calculation data (our results).

50

Figure 3.4. Coefficients of turbulent diffusion averaging by volume in central and peripheral parts of tubular apparatus of divergent-convergent design (Fig. 3. Ib) Peripheral - [-'-'7, central part - [ - - - - ' 1, by volume - T I .

0

-c

I

04 0,Ol

0.02

0,03

0.04

0,05 r. rn

Figure 3.5. Profiles of absolute values of rate in reaction zone sections EF, GH and PQ at y=45'(seeFig. 3.lb) Comparison of turbulent diffusion coefficients Dt in various regions of tubular reactor showed (Table 3.2) that apparatus of divergent-convergent design provides in volume the field of Dt homogeneous enough (for comparison of characteristic times of diffusion 7- and chemical reaction

51

the order of Dt value is significant). Application of reactor of divergent-convergent design with divergent angle of opening y in interval 20'-45' is expedient under practical realization of fast chemical processes. Table 3.2. Turbulent diffusion coefficients D,.103m2/sec averaging by volume in various parts and divergentconvergent sections (Fig. 3.lb) of tubular turbulent apparatus at different angles y. Re = 2,5105, dd = 0,Oj m, v = 5 d s e c , p = 1000 kg/m3

zch

10'

17'

30'

45'

60'

75'

85'

In apparatus of constant diameter (of cylindrical construction) initial parameters of turbulence at input, in particular flow geometry and the way of reagents introduction into apparatus (see 1.3.1) significantly influence on reaction mixture turbulization degree. At that Dt decreases as moving off inlet part and reduces thereby the intensity of liquid medium mixing along apparatus length (Fig. 3.6a). The usage of reactor of divergent-convergent design is advisable (Fig. 3.lb) for increasing of flow turbulization degree and consequently reagents mixing efficiency (Fig. 3.6b). Divergentconvergent canal allows maintaining high values of turbulence parameters along whole length of tubular apparatus made of several divergent-convergent sections of strictly limited extension (Fig. 3.6b). In apparatus of such construction turbulence parameters are determined by turbulization appearing at the expense of canals' geometry and they are in order or higher than turbulence level creating as in tubular canals of constant diameter, so in stirred tank reactors even under application of very effective mechanical mixers. In spite of divergent angle of opening y the ratio of divergent diameter to convergent diameter Q 1 d, and section extension L, / Q are key parameters allowing optimization of effectiveness of turbulent mixing in apparatus of divergent-convergent design. AS a rule the time of micromixing is limiting time of mixing under polymer synthesis fast process due to high viscosity flows use. The maximum of turbulence kinetic energy density dissipation E meets criterion of micromixing characteristic time minimum according with (1.18). That is why firstly it is necessary to study dependence of average value of turbulence kinetic energy density dissipation on ratio of reactor's geometrical sizes (Q / a,L, / Q). According with carried out calculations at fixed values of y and Q / & ratio section length L, increase leads to the fact that turbulence kinetic energy density dissipation E increases at first and then decreases. The analogous picture is observed at & / d, ratio change. This testifies to the presence of maximum point of average turbulence kinetic energy density dissipation E at some ratio of reaction zone geometrical sizes of divergent-convergent design. This ratio of sizes is optimal. Optimal ratio of geometrical sizes of tubular apparatus practically doesn't depend on angle y that simplifies the problem.

52

The criterion for optimization of geometrical sizes ratio of tubular turbulent apparatus of divergent-convergent design is the maximum of average value of turbulence kinetic energy density dissipation E , and parameters of optimization are the ratio of divergent-to-convergent diameters dd / de and the ratio of section length-to-divergent diameter L, / 4. Figure 3.7 illustrates the dependence of average value of turbulence kinetic energy density dissipation E on ratio of reaction zone geometrical sizes at y = 45'. The point of maximum M is indicated In Fig. 3.7 that corresponds to optimal parameters of reaction zone of divergent-convergent design: divergent-to-convergent diameters ratio dd / de = 1,6 and the ratio of section length-to-divergent diameter L, / & = 1,7. Optimal parameters are kept practically constant under angle y changing from 30' to 85'. Thus, tubular turbulent apparatus of divergent-convergent design is able to provide in any chemical conditions and mass-transfer processes homogeneous conditions at high turbulent mixing level appearing only at the expense of canal geometry without application of mobile inside facilities. In the view of intensification of turbulent mixing in apparatus of divergent-convergent design the optimal ratios are dd/d, = 1,6, Lddd = 1,7 and angle of divergent opening y is in interval 20'-45'. Design simplicity and high reliability and also mass-exchange effectiveness make this type of apparatus very perspective for carrying out of fast chemical processes. In the case of the last ones it is very difficult to provide criterion Da c 1 (zmix< T&) and to which very high safety requirements are made. Low convective rate of liquid flows in peripheral parts of tubular turbulent apparatus of divergentconvergent design (Fig. 3.5) and conical extension in apparatus of cylindrical construction (see 1.3.1) increase effectiveness of turbulent mixing and thereby reduce z&. However the last fact would influence on distribution of reagents residence times in reaction zone and its shape determines fast processes character proceeding and quality of resulted polymer products.

53

D

Figure 3.6. Distribution of turbulent diffusion coefficient (Dt) by reaction zone volume of cylindrical (a) and divergent-convergent (b) designs (Re =2.105, dd =0,05 m, V = 4 d s e c , p = 1000 kg/m3).

54

k E,

m2/sec3

\

Figure 3.7. Dependence of turbulence kinetic energy density dissipation sizes ratio.

E

on reactors geometrical

3.1.3. Distribution of reagents residence times in apparatus of cylindrical and divergentconvergent types The method of reception of curves of response to indicator introduction (see 2.2.8) is useful for studying of reagents residence times distribution in reaction zone and estimation of hydrodynamic regime of tubular turbulent apparatus operation with different canal's geometry (cylindrical and divergent-convergent) and reagents introduction way (coaxial and radial). For confirmation of adequacy of chosen technique for experimental determination of distributions of reagents residence times in tubular turbulent apparatus numerical calculation of three-dimensional turbulent liquid flow with the use of CFD soft PHOENICS was carried out. Axially symmetric nonswirling turbulent flow of continuous incompressible Newton's fluid was considered. In this case the generalized equation of substance (mass, impulse, heat, turbulence kinetic energy) transfer is: a ( p 0 ) /a7 + di4pu0) = di4pgrada + F (3.13) Here u - rate vector; F = f(O), - dependent variable indicating impulse of mass, enthalpy, turbulence kinetic energy units. Moreover field of indicator movement speed satisfies the law of conservation of mass (continuity equation) of the following form: (3.14) ap /a7 + di4pu) = 0 Turbulent stresses are determined by standard and modified K-E models of turbulence. For determination of flow field near the wall the method of parietal functions was used. Rates fields and pressures were corrected in the course of calculation according with SIMPLE-C algorithm. Diffusion of indicator was modeled as transmission of scalar introduced in less space of time in comparison with its average residence time in apparatus at fixed rate fields and turbulence parameters. Theoretical response curves were obtained, they characterized the flow structure and distribution of reagents residence times in tubular turbulent apparatus. Experimental and calculated

55

functions of reagents residence times distribution were correlated with coefficient 0,95 (Fig. 3.8), that confirmed the obtained experimental data adequacy. C

C

Figure 3.8. Distribution curves of reagents residence times in reaction zone. Apparatus V (a); VI (b), w = 130 cm3/sec (Table 2.2). Line - calculation; points - experiment. According with response curves (Fig. 3.8) tubular turbulent apparatus relate to apparatus of intermediate type, i.e. there is a deviation from plug-flow regime firstly at the expense of reverse current. The degree of the last one is determined by longitudinal mixing coefficient E. Distribution curves of reagents residence times are converged with mixing Z, to the low values region while volumetric speed of reagents w increases (Fig. 3.9) for all considered tubular turbulent apparatus constructions (Table 2.2). For description of hydrodynamic structure of reaction mixture flow in apparatus of cylindrical and divergent-convergent designs diffusion and cellular models were used (see 1.1.3). Values n + 1 (Fig. 3.10) and Bo + 0 (Fig. 3.11) correspond to ideal mixing regime and when displacement conditions are reached - n + m (Fig. 3.10) and Bo + 00 (Fig. 3.11).

56

0 .B

0.4

0 0

2

4

6

8

12

10

14

16,sec

18

Figure 3.9. The influence of reaction mixture movement on distribution of reagents residence times in reactor of cylindrical type. Apparatus I11 (Table 2.2); w [cm3/sec]: 130 (1); 91 (2);62 (3); 36 (4). N

I

0

6

0.5

1

1.5

2

2,5

3

Figure 3.10. Differential curves of distribution of reagents residence times in reaction zone (cellular model), n = 1 (1); 3 (2); 9 (3); 15 (4);21 (5); 30 (6).

c

I

6

Figure 3.11. Differential curves of distribution of reagents residence times in reaction zone (diffusion model), Bo = 1 (1); 6 (2); 18 (3); 30 (4); 42 (5); 60 (6). The inverse problem of response curves was solved for calculation of numerical values of Bodenstain's criterion Bo. For this purpose experimentally obtained differential response curves were approximated by dependences of cellular (1.9) and diffusion (1.11) models (Fig. 3.12). The biggest disarrangement between values calculated by (1.9) and (1.1 1) and corresponding experimental response curves didn't exceed 15% for all studied apparatus, and average disarrangement was 7%. The influence of canal geometry on hydrodynamic regime of operation is characteristic for apparatus with coaxial reagents introduction (Fig. 2.7a,b). For apparatus of divergent-convergent design (Fig. 2.7b) reduction of ratio of section length-to-divergent diameter L, / dd (apparatus V, VI), of diameter dd (apparatus VI, VI11 (Table 3.3)) and also the increase of canals profiling degree, Le. ratio dd / d, (apparatus VI, VII) approach the liquid reagents structure to plug-flow regime that is determined by low values of criterion Bo (Table 3.3, Fig. 3.13). Thus, divergent-convergent design sections function as statistical turbulization devices increasing turbulent diffusion coefficient Dt (Fig. 3.6) and longitudinal mixing rate (Fig. 3.14).

58

m

0

OS

2

1

Figure 3.12. Experimental (points) and calculated (Bo = 57 (firm line), n = 28 (dotted line), w = 91 cm3/sec) differential curves of reagents residence times distribution in reaction zone. Apparatus I11 (Table 2.2).

BO

60

-

40

-

I

3

20

60

100

140

w,crn'lisec

Figure 3.13. Dependence of Bodenstain's criterion on liquid flows' speed w in reactor of divergentconvergent design (v-VIII, see Table 2.2,3.3).

59

I

IV

40

I

/w

20

I

0

-

v 100

0

4 v

In

.

*I

w, cm’isec

I

200

Figure 3.14. Dependence of liquid flows longitudinal mixing coefficient E on apparatus geometry (I-VIII, see Table 2.2, 3.3) and reaction mixture volumetric rate w. Table 3.3. Parameters of hydrodynamic structure of reaction mixture movement in dependence on reactors

‘WI = 24 cm3/sec; w2 = 80 cm3/sec correspondingly (Fig. 2 . 7 , ~ ) .

(w1,

wz - volumetric rate of central and axial flows

Apparatus of cylindrical type with coaxial reagents introduction (Fig. 2.7a) with dd 2 0,03111 (apparatus 1-111) is characterized by high values of criterion Bo (Table 3.3, Fig. 3.15) and consequently by higher in comparison with apparatus of divergent-convergent design approximation degree of reagents flow structure to plug-flow regime due to low longitudinal mixing rate (Fig. 3.14). In consequence there is narrow distribution of reagents residence times in reaction zone. As flows structures in apparatus of cylindrical design with dd 10,03m are practically equal (Bo = 50) so there is no necessity in application of turbulent reactors of large diameters in industry

60

for polymer synthesis that allows creation of compact highly productive technologies. It is characteristic that in apparatus of any structure with coaxial reagents introduction (Fig. 2.7a, b) the volumetric rate of liquids flows (apparatus productivity) practically doesn't influence on criterion Bo (Fig. 3.13, 3.15). However, increase of longitudinal mixing rate occurs in this case.

04 20

60

140

100

,".crn'/ree

Figure 3.15. Dependence of Bodenstain's criterion Bo on reaction mixture movement rate w in apparatus of cylindrical design (I-III, Table 2.2, 3.3). For apparatus with redial flows introduction (Fig. 2.7b) not only canal geometry but also the ratio of rates of initial reagents introduction w1 and w2 influence on hydrodynamic regime of their operation (Fig. 3.16). At fixed speed of central flow w1 increase of w2 leads to the rise of criterion Bo and approximate the operating regime of apparatus of this construction to plug-flow regime. At high rates of central flow w1 the leveling of radial flow rate w2 influence on hydrodynamic structure of reaction mixture movement occurs (Fig. 3.16). BO

160 1

5

120-

b

A

3

3

e

80 4 8

1

18

28

38

48 WL

rrlhec

Figure 3.16. Dependence of Bodenstain's criterion Bo on the rate of central reagents flow w2 = 41 (1); 47 (2); 57 (3); 67 (4);80 (5) cm3/sec. Apparatus IV (see Table 2.2, Fig. 2 . 7 ~ ) .

w1

at

The observed dependence of tubular turbulent apparatus work on the way of reagents introduction is explained by the fact that in this case the formation of various reaction fronts macrostructures in mixing zone is possible, in particular planar front (P), torch (T) and drift (Z) etc. (Fig. 2 . 7 ~ (see ) 1.3.1). At that the macrostructure in the form of planar reaction front corresponds to

61

quasi-plug-flow regime in turbulent flows determining the optimal conditions of fast chemical processes realization under polymers synthesis at quasi-isothermal conditions, that is very favorable for production. The increase of radial flow rate w2 leads to the transition between reaction fronts in following order - Z + ll + T (Table 3.4). the rise of coaxial rate w1 leads to the inversion of order of transitions between macrostructures (Table 3.4), criterion Bo approach to constant value about Bo = 120 at that. Table 3.4. Conditions of characteristic reaction fronts macrostructures formation in dependence on the ratio of reagents introduction rates.

It is important that in the case of formation in reagents mixing zone the only possible in given experimental conditions macrostructure in the form of drift in turbulent apparatus of cylindrical design with radial reagents introduction (Fig. 2 . 7 ~ )the influence of reagents movement rate on reactor's operating regime would not be observed. However, reaction front of Z type is not desirable under fast processes proceeding in polymers synthesis due to reagents delay and residence time increase z*. Reaction mixture movement character in reaction zone is characterized also by parameter Or (see 1.1.3).For its determination average-integral residence time zr was used (1.12): it is a ratio of functions' integrals ZiCi = f(q) M Ci = f(Zi) (Fig. 3.17). Nil@,mlel

hc+105.lsec mle)A)

12.

8.

4.

0,

-.

- 1

4

Figure 3.17. Dependence of Ci(ziCi) on zi for determination of average-integral residence time 7,.1 - Ci - Ti; 2 - ZiCi - Ti. Apparatus III,w = 130 cm3/sec (Table 2.2.)

62

Tubular turbulent apparatus of cylindrical design with diameter & I 0 , 0 3 m (apparatus I-III, Table 2.2) have 0,I 1 ( ~ a 6 n 3.5) . that is characteristic for plug-flow regime. Obviously it relates to the fact that in apparatus of such construction turbulent mixing level (Fig. 3.6a) and consequently the gradient of turbulent pulsations directed opposite to main flow reduce along apparatus axis x as moving off regents introduction zone. This determines narrow distribution of reagents residence times in reaction zone. At the same time, tubular turbulent apparatus of divergent-convergent design (apparatus V-VIII) are characterized by zones of product delay (0, > 1, Table 3.5) increasing reagents residence times in reaction zone. Table 3.5.

Indicator delay

Thus, the influence of reaction zone geometry and reagents introduction way on hydrodynamic reaction mixture movement structure and reactors operating mode was revealed with the help of solution of response curves inverse problems. At coaxial reagents introduction the hydrodynamic regime of tubular turbulent apparatus operation is determined by reaction zone geometry. At radial reagents introduction in one and the same apparatus in dependence on the ratio of reagents introduction rates the formation of modes with various approximation degree both to idealized mixing and displacement models is possible. The last fact is impossible in the case of existing reactors of applied chemistry. Noticeable differences between tubular turbulent apparatus of cylindrical and divergentconvergent design determine various spheres of their industrial application under fact processes realization in polymer synthesis. Tubular turbulent apparatus of cylindrical with dd < 0,03m and divergent-convergent with dd 5 0,03m; L, / dd = 2a3, dd / d, = 1,6+3 designs are characterized by intensive longitudinal mixing and presence of zones of reagents circulation that approach them to apparatus functioning in ideal mixing regime. In cylindrical apparatus with diameter dd < 0,03m the rate of longitudinal mixing is comparatively not high (E 5 4.10" m2/sec) and consequently there are practically no zones of reagents delay or slip (0, = l), i.e. they work in quasi-plug-flow mode with narrow reagents residence times distribution. Reaction zone of cylindrical design is advisably to be used for realization of extra-fast chemical processes with characteristic reaction time zch I o , o 1 in quasi-plug-flow mode when the process to 100% proceeds at places of reagents introduction and where the maximum level of turbulent mixing is reached (Fig. 3.6a). Divergent-convergent construction is effective under realization of fast exothermal reactions, at that intensive longitudinal mixing of liquid flows allows creation in reaction zone of quasi-isothermal (isothermal in any cross section of flow) regime. Intensive turbulent mixing allows recommendation to use tubular turbulent apparatus of divergent-convergent design as prereactors in processes where fast and slow processes proceed simultaneously to remove diffusion restrictions at fast stages of "gross" processes. High level of turbulent mixing and its constancy over the whole reaction zone volume (Fig. 3.6b) and consequently high approximation degree to ideal mixing regime (Fig. 3.13) determine prospects of application of apparatus of divergent-convergent design for intensification of mass-exchange processes both in homogeneous and heterogeneous flows and also for intensification of convective heat-exchange at external thermostating of fast polymerization reactions.

63

According with obtained results of studying of reagents residence times distribution in tubular turbulent apparatus we may conclude: Tubular turbulent apparatus in dependence on canal geometry are characterized by various approximation degree to reactors of ideal mixing and displacement; Compact tubular turbulent apparatus of cylindrical design with dd 1 0,03m are characterized by comparatively low rate of longitudinal mixing E 54.10” m2/sec, relative flows residence time 0,s 1, narrow reagents residence times distribution, that allows realization of fast chemical reactions with Tch I0,01 sec in quasi-plug-flow mode; Tubular turbulent reactors of divergent-convergent design at L, / dd and dd decreasing and dd / de increasing approach apparatus of ideal mixing with broad reagents residence time distribution due to high rate of longitudinal mixing. Apparatus of this design is advisably to apply for realization of fast exothermal reactions in quasi-isothermal conditions. At the same time there are no data in literature about limitations of tubular turbulent apparatus functioning effectiveness in relation to liquid flows viscosity that caused significant difficulties at implementation and effective exploitation of this apparatus in polymer production. In particular in polymers synthesis in solutions, for example in CBR production at stages of butyl rubber chlorination by chlorine in solution, washing out and neutralization of highly viscous polymerizate, etc. This fact determines the expediency of studying of the influence of liquid flows viscosity on effectiveness of turbulent mixing in tubular canals. It is also important to find the ways of the following situation realization: when high viscosity of applied polymers solutions and other high-viscosity mediums doesn’t play any role under reagents mixing and fast polymerization processes proceeding. 3.1.4. Fast processes proceeding in high-viscosity mediums 3.1.4.1. Self-similarregime o f reaction medium flow

The main problem of polymer products production in solution, in particular ethylene propylene and chlorobutyl rubbers, is viscosity rise on one-three levels while polymer concentration in solution increases. High viscosities of polymers solutions as a rule do not provide intensive mixing of reaction mixture because values of Reynolds criterion (Re) and consequently turbulent diffusion significantly decrease. At that ineffective molecular diffusion is the most possible way of mixing. That is why, for example under polymerization, decomposition and washing out of catalyst, and also under polymer stabilization before its isolation they work with solutions diluted enough (not higher than 10-13 mass %) and it is a substantial disadvantage of acting productions. With monomer concentration decrease in reaction mixture polymerization rate and production efficiency reduce. At the same time it is known (see 1.3.1) that in high-turbulent flows viscosity doesn’t influence on base medium volume flow characteristics because of self-similar liquid flow formation in relation to criterion Re and viscosity. The value of criterion Re over which self-similar region is observed is determined by reaction zone geometry. They may expect that self-similar regime will begin to form at significantly lower Reynolds number because in comparison with cylindrical canal in divergent-convergent design of turbulent apparatus at the same Reynolds values significantly higher turbulization degree of flow is achieved (Fig. 3.6). Numerical solution of equation of liquid flows turbulent movement with the use of IC-& turbulence model (3.1)-(3.12) confirmed this suggestion. In divergent-convergent canal self-similar regime comes at Recr= 800/f. Parameter f is the function of angle of divergent opening y and its values can be found by graphical dependence (Fig. 3.18) or by formula approximating this dependence: f = 0,117 +O,O49y 0,0012$ + 1,374.105f’ 5,9.108y4 (3.15)

-

-

64

In particular in interval of divergent opening angles y 30'-80' self-similar regime is observed at Re 2 Recr = 950f50 (Fig. 3.18). At the same time in apparatus of cylindrical design self-similar regime of reaction mixture flow is observed at Re > lo', i.e. is higher in four orders. Re,,lllJ

f

LO

0

0

40

80

y, degrees

Figure 3.18. Dependence of critical value of Reynolds criterion (Recr) and parameter f on divergent opening angle (y) in reactor of divergent-convergent type. Thus, in self-similar regime all values determining reagents turbulent mixing when medium homogenization is limited by exchange processes between large turbulent flows don't depend on flows viscosities (self-similar flow in relation to Re). This fact constricts the circle of values determining properties of reagents turbulent mixing in tubular apparatus of jet type. Only three values characterizing large-scale movements remain: medium density p, apparatus diameter d and linear rate of flows movement V. The level of liquid flows turbulence depends on these three values at conditions of its independence on viscosity. With the help of these three values they may constitute the following single combinations with corresponding dimensions for average values of turbulence kinetic energy density K and its dissipation, turbulent diffusion coefficient Dt and hydraulic resistance Ap:

-

(3.16) IC,, V2; E ~ V V3/d; D t a v -V'd; Ap -p.V 2 In particular, the last fact can be illustrated by the example of pressure loss under liquid flows movement in cylindrical canals, where at laminar flow (Re 22300) Ap depends on Re and doesn't depend on canal's wall roughness [3, 1351:

3 2 p .L.v2 (3.17) Re d ' and at turbulent movement (Re = (4-100).103) in self-similar region of flow in relation to Re, Ap stops to depend on liquid movement character and depends only on wall roughness [3, 1351: A6=-

(3.18)

Here 6 - relative roughness of wall equal to ratio of average height of roughness on inside surface of tubing (absolute roughness) to tubing diameter.

65

We succeeded in finding of numerical coefficients in (3.16) by data file handling received by solution of equations of continuum turbulent movement with the use of IC-& turbulence model (3.1)(3.12) and by method of finite elements on irregular computational grid. The average value of turbulence kinetic energy density: IC = fcf 2v,2 (3.19) The average rate of its dissipation:

fEf 3v,3

(3.20)

&=-

dC

The average value of turbulent diffusion coefficient:

DT =

0,09f:fVcdc

(3.21) fE where coefficients fc ki fE are determined by geometric parameters of reaction zone & / & and Ls/ &: L d L Ls 2 t 8,64.104.(+3 + 0,0784 x)+ fc= -0,074 + 0,0124 A )- 8,74*103*( -) dd dd dd dC

dd )2 -3,224 dd )2 *(Ls ) dd )*(Ls )2 - 0,0224 dd )*(3 ) - 1,31*103 *(t0,021*(dC dc dd dc dd dc dd (3.22)

fE = -0,138 t 0,2264 d ) dC

dd )3 4- 0,034 A L )-4,95*103**(dd )*(-)Ls dd )2 + 0,0194 - 0,1164 dC

dd

dC

dc

dd

dd )2-9,62*10'34L )2+3,22.103*( dd )-(Ls )2 1,93*10"(A )*(dd dc dd dc dd (3.23) The possibility of self-similar regime formation in reactor of divergent-convergent design allows reception of acceptable for engineering calculations formulas of mixing characteristic times by substitution of (3.19)-(3.21) to (1.17)-(1.19). Characteristic time of turbulent mixing: 11,112fE (3.24) z = turb fcZfVcdc In this case characteristic time of micro-mixing (mixing at the expense of molecular diffusion) can be calculated by ratio: ..

z

micro

(3.25)

Characteristic time of meso-mixing (mixing at the expense of exchange between large turbulent flows and being inside of them little flows) is calculated by ratio: (3.26)

G6

Formulas (3.24)-(3.26) are suitable for calculation of turbulent mixing characteristics in fOllOWing ratios intervals: dd / dc = 1,2+2,5;Ls /dd = 0,5+3,5; Ls / dd > (l-dd/ dc)cta. Comparison of mixing characteristic times calculated by (3.24)-(3.26) with chemical reaction characteristic time Tch or liquid flows residence time in apparatus Tr allows calculation of optimal construction of tubular turbulent apparatus for both fast chemical reactions realization and flows mixing with the aim of their homogenization. Formation of optimal hydrodynamic regime in reactor of divergent-convergent design causes possibility of polymer products concentration increase in solution under rubbers and thermoplastics synthesis and also under operation with their high-viscosity solutions up to the stage of resulted products isolation including fast chemical reactions realization. There is an exigency of transition of a number of polymer productions in solutions to the resources- and energy-saving highly productive technologies of heightened environmental safety with wide use of flowing compact tubular turbulent apparatus of divergent-convergent design. Thus, the possibility of realization of self-similar flow regime in tubular turbulent apparatus of divergent-convergent design at comparatively low reagents movement linear rates enlarges the spheres of their industrial usage for the work with high-viscosity mediums and also allows reception of equations for calculation of average values of turbulent diffusion coefficient Dt, turbulence kinetic energy density K, its dissipation E and characteristic times of flows mixing at various scales. With the aim of optimization of fast processes realization conditions under polymer synthesis it is advisable to fix the correlation between geometrical sizes of tubular turbulent apparatus of divergent-convergent design, its operation dynamics, liquid flows physical parameters and average values of turbulent mixing characteristic time in reaction zone.

3.1.4.2. Outimization o f characteristics o f reagents turbulent mixing in self-similar regime Under fast chemical reactions realization for the exception of diffusion impediments the following condition should be met: Tfix = Tturb I Tch, where Tch is chemical reaction characteristic time (under polymerization - Tch = 1 / kdeac.;under fast processes with participation of lowmolecular reagents - 7ch = 1 / kc""). At the same time, high level of reagents mixing and consequently optimal conditions for chemical processes realization are determined by ratio of characteristic times of liquid flows mixing on macro- and micro-levels Tturb > Tficro. If medium viscosity is increased according with (3.24) and (3.25) fast chemical and correspondingly mass-exchange physical processes gradually pass to the region limited by micromixing, i.e. Tturb > Zficro condition is not fulfilled (Fig. 3.19). That is why when one works with high-viscosity polymer solutions in stirred tank reactors and also in tubular turbulent apparatus with low rate of flows reagents mixing occurs due to ineffective molecular diffusion. Application of reactor of divergent-convergent design under rise of reagents movement linear rate allows transformation of reactor's work to the turbulent regime when mixing time will be limited by effective large-scale turbulent exchange (Fig. 3.19). It is also obvious from Figure 3.19 that under the rise of reaction mixture flows rate up to 4 d s e c significant decrease of characteristic times of micro- and macro-mixing is observed.

0

1

2

3

4

5

6

V , h k c

8

Figure 3.19. The values of characteristic times of turbulent mixing Zturb (l), micro-mixing Zmicro (3-61, and meso-mixing zms0(2) in dependence on reaction mixture movement rate V. The values of dynamic viscosity: 0,001 (3), 1 (4),10 (5),50 (6) Pa.sec. d = 0,025m, p = 1000 kg/m3, V, = 4 dsec. Under further increase of V the change of mixing times values is not so significant. Thus, one can speak about critical value of liquid flows rate in tubular canal V,, = 4 d s e c ; it is possible not to exceed it in real conditions. However, at the expense of increase of reaction mixture movement rate (apparatus productivity) turbulence kinetic energy density is increased (Fig. 3.20) and by this intensification of diffusion processes occurs. Optimal characteristic times of mixing is also determined by interval of divergent opening angles y higher than 30' (Fig. 3.21). At the same time, for example for liquid flows with viscosity 50 Pa.sec it is impossible to reach optimal conditions for mass-exchange processes proceeding Zturb > Zmicro at technologically acceptable angles y (Fig. 3.21). In this case one should increase reaction mixture movement rate V (Fig. 3.19). Characteristic times of turbulent, micro- and meso-mixing are the important characteristics determining the possibility of tubular turbulent apparatus application for concrete process of chemical technology and also its geometrical parameters. For example under fast chemical reaction when process practically completely proceeds at reagents delivery points significant role is played by numerical values of meso-mixing characteristic time znms0 - exchange between turbulent flows large enough and situated inside of them little ones. In the case of emulsification or agglomeration processes the average size of drops (particles) of disperse phase depends on flows mixing on microlevel and is determined by characteristic time of nicro-mixing Zdcro. The time of reaching of required quality of mixing (mixture residence time in apparatus) should be comparable with characteristic time of large-scale turbulent mixing Zturb when tubular turbulent apparatus of divergent-convergent design are used for homogenization of liquid flows. In general case, for optimal proceeding of processes limiting by mass-exchange in turbulent flows confined by impenetrable grid the following ratio should be met: Z,I~> Zturb > Znmso > Znlicro.

68

2

0

*

6

4

V. misec lo

Figure 3.20. The values of average turbulence kinetic energy density K,, along volume of reaction zone of divergent-convergent design (y = 45’) in self-similar regime in dependence on reaction mixture movement rate V. tmix, sec

1 3

1

0

10

20

30

40

50

60 g, degrees

Figure 3.21. The values of characteristic times of turbulent mixing ‘&b (I), micro-mixing ‘Cmiwo (3-6), and meso-mixing zms0( 2 ) in self-similar regime in dependence on divergent opening angle y. The values of dynamic viscosity: 0,001 (3), 1 (4), 10 (3,50 (6) Paesec. d = 0,025m, p = 1000 kg/m3, V, = 4 d s e c .

69

According with (3.24)-(3.26) the magnitudes determining values of characteristic times of mixing are linear rate of liquid flows movement V, apparatus diameter d, divergent opening angle y, and kinematic viscosity v for micro-mixing. Variation of diameter and reagents movement linear rate is the only accessible method of effect the homogeneity of reaction mixture in reaction zone of divergent-convergent design (Fig. 3.22-3.24). It is obvious that optimal ratio Tturb > Zms0 > Zmicro is nearly always fulfilled.

Figure 3.22. The dependence of turbulent mixing time Tturb on apparatus diameter d, and linear rate of reaction mixture movement Vc. y = 45', dd / d, = 2, L, / dd = 3. Characteristics of liquid flows, in particular density and viscosity significantly influence on reagents mixing on micro-level (Fig. 3.25). Rise of viscosity and reduce of density of reagents introducing into tubular turbulent apparatus may lead to limitation of reaction mixture homogeneity by ineffective molecular diffusion that happens when one works with polymers solutions very often. Analysis of dependences (3.24)-(3.26) allows suggestion of criterion of violation of self-similar regime of reaction mixture flow, i.e. condition of reduction of efficiency of operation of tubular turbulent apparatus of divergent-convergent design with medium viscosity increase:

P 0,412Pf ;fvc

->

P

f:d:

(3.27)

70

&io3, m

Figure 3.23. Dependence of characteristic time of meso-mixing zmsoon apparatus diameter d, and linear rate of reaction mixture movement V,. y = 45', da / d, = 2, L, / dd = 3. In this case tubular apparatus operation may be optimized at the expense of flows linear rate rise in accordance with ratio zturb - 1 1 v, zmso- 1 1 v, Zficro - 1 / Vlf5that also allows significant rise of process productivity w because w - V. The increase of linear rate of reaction mixture movement provides optimal values of characteristic times of liquid flows mixing (Fig. 3.22-3.24), turbulent diffusion coefficient (Fig. 3.26) and turbulence kinetic energy density dissipation (Fig. 3.27). High level of tubular turbulent apparatus application by dynamic characteristics of their work in this case is obviously pressure fall at apparatus ends in accordance with Ap - V2 and the low level is D, I m2/sec (transient condition).

dsec

d,103, m

Figure 3.24. Dependence of characteristic time of micro-mixing 7"icroon apparatus diameter de and linear rate of reaction mixture movement V,. y = 45', p = 1000 kg/m3, p = 1 mPasec, dd / d, = 2, Ls 1 dd = 3.

71

Figure 3.25. Dependence of characteristic time of micro-mixing r,,,icroon density and viscosity of reaction mixture. y = 45', & = 0,025 m, Q I & =2, LsI Q = 3, V, = 4 d s e c .

Figure 3.26. Dependence of turbulent diffusion coefficient D, on apparatus diameter dc and linear rate of reaction mixture movement v,. y = 45O,Q / & =2, L~/ Q = 3.

72

Figure 3.27. Dependence of turbulence kinetic energy density dissipation E on apparatus diameter d, and linear rate of reaction mixture movement V,. y = 45', Q / & =2, L, / Q = 3. Decreasing of reactor's diameter leads to reduction of mixing characteristic time and this is the key to the realization of fast processes in optimal conditions. However it leads to the reduction of effective turbulent diffusion coefficient Dt (Fig. 3.26). In particular values of Dt determine the low limit of tubular turbulent apparatus application possibility by their geometrical parameters under conditions of industrial production. Calculations show that at d, < 0,023 m, V, = 4 d s e c and y = 45' the diffusion coefficient is D < m2/sec that is characteristic for transient condition of liquid flows in cylindrical canals. Application of reactor of lower diameter leads to increase of average values of turbulence kinetic energy density dissipation E (Fig. 3.27). The maximum value of E determines intensity of liquid flows mixing on micro-level (kolmogorovb scale) that provides appearance of small-scale shear deformations and consequently obtaining of thin-dispersive emulsions and suspensions. In this case decrease of diameter of tubular turbulent apparatus of divergent-convergent design and increase of reaction mixture movement linear rate are adequate to the increase of number of revolutions and mechanical mixer fans diameters in stirred tank reactor. Concerning the length of mixing zone in turbulent flows calculations of characteristic times of mixing and chemical reaction allow derivation of criterion of possibility of fast processes realization in chemical technology in tubular turbulent apparatus: under polymer synthesis (polymerization): V L 2 L = v z =A (3.28) ch c ch k d under low-molecular products synthesis in flow:

LIL under flows homogenization:

= vc zch = ch

V

C

k[C]" -

(3.29)

13

ll,lPfE

L2L

mix

= vc 2turb =

fc2fdc

If one assume that 1 = d then one may convert: 11,ld cfE L2 fLf

(3.30)

(3.31)

C

For process not fast enough possibility of tubular turbulent apparatus use is limited by its length large enough at not high values of reaction rate constant. At the same time under carrying out of mass-exchange physical processes in particular flows mixing with the aim of homogenization required extension of mixing zone and consequently apparatus length depend only on its canal's geometry. This provides wide possibilities of optimization of practically any process where it is necessary to create homogeneous filed of reagents concentrations. Analysis of (3.28)-(3.30) shows that fast chemical reactions proceeding in optimal conditions, i.e. in quasi-plug-flow mode in turbulent flows the following condition should be fulfilled Lch 1 Lfix:

vc k[C]"-l

vc

ll,l/LfE 2kd f:fdc

(-)

(3.32)

According with (3.32) it is not difficult to determine critical values of chemical reaction rates constants which are necessary for fast processes realization in the absence of diffusion limitations. In Figure 3.28 the dependence of critical values of rate constant of low-molecular reaction of the second order on reaction mixture movement linear rate V in tubular turbulent apparatus and also on its construction is presented as an example. Increase of V and reaction zone diameter d allows carrying out of chemical reactions in optimal conditions with values of rates constants high enough. In particular, at technically acceptable values of d and V chemical reaction proceeding in the absence of diffusion resistances is limited by the value of constant of low-molecular compounds reaction rate - k I10 Ymole.sec. Under proceeding of more fast liquid-phase chemical reactions with k >> 10 Vmole.sec in turbulent flows with application of compact tubular turbulent apparatus of divergent-convergent design for reaching of ratio 2ch 2 2turb it is necessary to increase linear rate of flows v, to reduce d and to decrease corresponding concentrations of introducing into reaction zone reagents sharply. Thus, when changing the geometry (design) of tubular turbulent apparatus of divergentconvergent construction, dynamics of its operation and also physical parameters of reagents liquid flows one may optimize the values of turbulent mixing characteristics in accordance with proceeding process specificity limited by mass-exchenge. There are intervals of values of diameter of tubular turbulent apparatus of divergent-convergent design and of liquid flows movement linear rate at which conditions for diffusion limitations taking off for fast chemical reaction proceeding are created. In accordance with process character (kinetic parameters, physical characteristics of liquid flows, etc.) regularities obtained in the work allow scientifically-grounded selection of optimal conditions for its carrying out (reaction zone geometry, dynamic regimes, etc.). In spite of processes occurring in homogeneous mediums chemical reactions proceeding at interface in systems "liquid-liquid'' and "liquid-gas'' when process rate is limited by substance diffusion rate through dividing surface are prevalent in various branches of chemical, petrochemical and others industries. In this case the supreme problem is to lower influence if limitations for reagents diffusion, in particular at the expense of homogeneous dispersions with developed phase contact surface.

74

k, Ilmole sec

3

0

4

8

Figure 3.28. Critical values of rate constant of reaction of second order proceeding in kinetic region in dependence on reaction mixture movement rate V and tubular turbulent apparatus diameter d, = 0,025 (l),0,02 (2), 0,Ol (3) m at C = 1 mole/l.

3.2. Multi-phase systems 3.2.1. Mathematical modelling Irregular or multi-phase reaction systems are characterized by the presence of macroscopic (in relation to molecular) heterogeneities. Numerical calculation of hydrodynamics of such flows is very complicated. There are two principally different approaches to their description: Eiler's approach taking into account interphase interaction (model of interpenetrative continuums) and Lagrang's approach consisting of integration upon trajectory of discrete particles (drops, bubbles, etc.) movement. Lagrang's approach is unacceptable for studying of multi-component systems because real systems contain significant amount of particles of discrete systems. According with Eiler's approach two-phase flow is presented as two interpenetrative continuums, Le. every phase is considered as quasi-continuous phase but under interphase interaction conditions. When carrying out mathematical modelling one supposes following assumptions: 1) sizes of particles of disperse phase are in many times larger than molecular-kinetics; 2 ) sizes of disperse particles are in many times smaller than distances on which averaged or macroscopic parameters of mixtures or phases are significantly changed (except separate zones that are considered as break surface), i.e. heterogeneities sizes under multi-phase flows are significantly smaller than sizes of divergent-convergent canals (L and dd). The first assumption allows application for calculations of classical ideas and mechanical equations of continuous homogeneous mediums for description of processes in scales of heterogeneities themselves (drops, bubbles, etc.). The second assumption determines possibility of description of macroscopic processes in multi-phase mixture by methods of continuous medium mechanics with the help of averaged or macroscopic parameters.

75

The description of mechanics of continuous medium of multi-phase reaction mixtures accordingly with Eiler's approach is connected with introduction of conception of multispeed continuum and determination of interpenetrative motion composing disperse system. Multispeed continuum is totality of N continuum, each of them relates to its composing mixture (phase or component) and fills up one and the same volume. Density p, continuum rate and then other parameters were determined for each composing continuum in every point by usual method. Thus, in each volume point filling up by mixture N densities, N rates etc. are determined. Furthermore, parameters characterizing components mixture as a whole such as density and bulk mixture flow rate can be determined on the base of these magnitudes. A lot of problems of hydrodynamic and heat-mass-exchange including multi-phase flows defy analytical calculation and the only possibility of their theoretical analysis is reception of numerical solution. Methods of equations discretization on nonorthogonal curvilinear grids allowing tracing of calculating regions borders shape get a wide distribution for this aim. In the field of heat-masstransfer and hydrodynamic study CFD (Computational Fluid Dynamics) system is generally aacepted. It allows numerical prediction of gases and liquids movement character, heat-transfer, phase conversions, chemical reactions, etc. basing on classical laws of physics and chemistry. One of the most popular and widely used unique program package for solutions of such problems is PHOENICS (see 1.3.2) allowing investigation of wide diapason of physical processes including multi-phase flows in systems "liquid-liquid'' and "liquid-gas". This unique program package allows plotting of various geometry regions, and calculation grid by this may be curved and optimized for any geometrical surface, in particular for reaction zone designs of cylindrical (with conical extension) and divergent-convergent designs. For investigation of mediums movement in turbulent flows unique program package PHOENICS has more than twenty turbulence models, in particular standard IC-& model and its modifications and also different from it models, for example one based on length of mixing way, renormalized IC-o model, etc.. Correlations of experimental results introducing into slip model IPSA determine large selection of inter-phase interactions variations participating in wide region of Reynolds' numbers and allow modelling of movement of disperse phase particles of different shapes (spherical, ellipsoidal, etc.). Numerical solution of averaged by Reynolds equations of continuous medium movement with the use of unique program package PHOENICS were chosen for investigation of turbulent movement of flows in reaction zone of divergent-convergent design (Fig. 3.lb). Program package PHOENICS was used for numerical solution of generalized equations of substance transfer (mass, impulse, heat):

dbiPi6 i)

dz

where ai, pi, pi,

++ + di~ctipiI+ Oi - aipigradq) = ai Fi ,

(3.33)

+

u1 are volume fraction, density, transfer coefficient (viscosity) and vector of i-

phase rate, accordingly (1 - continuous, 2 - disperse phase). Non-stationary, convection and diffusion terms are in the left part of equation and source term in the right. Dependent variable Q signifies various values in particular component rate, reagents mass concentration, turbulence kinetic energy or turbulence scale. Equation (3.33) is added by continuity equations expressing law of conservation of mass:

+

dbiPi)

dt

+di4a,p, u i ) = O

(3.34)

1 1

Differential equations (3.33) and (3.34) are easily solved for stationary regime of axial symmetric non-swirling turbulent flow of incompressible Newton two-phase medium without consideration of inter-phase heat- and mass-transfer. That is why the force of inter-phase interaction

+

at the expense of friction presents source of E in (3.33). One assumes that all disperse inclusions 1

(drops, bubbles, etc.) have spherical shape.

16

The force of inter-phase interaction is connected with the rate of slipping by coefficient of inter-phase interaction:

+ +

F = f Vslip (3.35) Slipping rate is determined by difference between rates of continuous V1 and disperse V2 phases:

+ +

+

Vslip = V 1 - V 2 Coefficient of inter-phase interaction depends on density continuous medium and friction coefficient C :

(3.36) PI,

ri

f=O,SCFplQ V slip vr,

volumetric fraction

a1

of

(3.37)

where F - specific surface area of phase contact; vr - volume of reaction zone. Friction coefficient C depends on Re numbers: for Re > 100

C=-1 6 Re

for Re < 0,49 for 0.49 < Re < 100

2069 it is If Re >> 100 and We > 8 friction coefficient is C = 2,7 and at Re >> 100 and Re > we2v6 determined as C = We / 3 (We - Veber's criterion). Number Re for multi-phase flows was calculated as:

+

Re =

V slip d2 V

(3.38)

where d2 is the average diameter of particles of disperse phase; v - kinematic viscosity of twophase mixture. Veber's criterion is the measure of ratio of inertia to inter-phase tighting force and characterizes the constancy of this ratio in corresponding points of such flows:

dfv

We=-

(3.39)

The characteristic particularity of reactor of divergent-convergent design (Fig. 3. lb) is the appearance of circulation regions (Fig. 3.29), Le. when near to impenetrable wall under the action of pressure gradient the particles of disperse phase in flow begin to move opposite to direction of main reaction mixture flow. The following turbulence models may be used for numerical studying of particularities of such flows (see 1.3.2): standard k-E (KEMODL); modified k-E (KECHEN) and renormalized k-&(KERNG) turbulence models.

Figure 3.29. Vector field in reactor of divergent-convergent design (Q = 0,08 m, & = 0,05 m, LS = 0,27 m, V = 16 m, p1= 6,93 kg/m3, pz = 640 kg/m3, az= 0,2) Calculations made with the use of PHOENICS confirmed the opinion that standard k-E turbulence model actually gave overestimated values in particular of turbulence kinetic energy K. Significantly better results were obtained with the help of KECHEN and KERNG models (Fig. 3.30), that were worked out especially for separation flows with circulation zones (Fig. 3.29). K, m2Isec2

Figure 3.30. The change of turbulence kinetic energy density K along axis AC of apparatus of divergent-convergent design (Fig. 3.lb). Calculations by KEMODL (l), KECHEN (2) and KERNG (3) models. Under fast processes of polymer synthesis proceeding on the phase interface in reactor of divergent-convergent design it is necessary to be able to estimate the average diameter of drops (bubbles, etc.) of disperse phase and their distribution by sizes. In this case the values of rate of turbulence kinetic energy density dissipation E play a defining role in accordance with Kolmogorov's isotropic turbulence theory. At self-similar mode (Re > lo3) (see 3.1.4.1) for tubular apparatus of divergent-convergent design averaged by volume turbulent flow characteristics practically don't depend on density, but on

78

values of phases densities p1 and p2, their volumetric rates w1 and w2, apparatus diameter d, and middle-outlays mixture rate Vc. The combination for average value of rate of turbulence kinetic energy density dissipation of two-phase flow may be consisted in accordance with dimensions analysis when using significant values: (3.40) Eav (Vc, dc, PI, P2, ~ 1~ , 2 ) Numerical solution of (3.33) and (3.34) by geometrical mixing zone parameters, physical characteristics and dynamics of disperse systems with corresponding boundary condition allows obtaining of analytic solution of equation (3.40): - 0.3

-

av Here f and f E are calculated from (3.15) and (3.23). Maximum value of turbulence kinetic energy density dissipation determined as:

(3.41)

in this case is

(3.42) The biggest discrepancy between values calculated from (3.41), (3.42) and corresponding results of numerical calculations does not exceed 23%, and average discrepancy is about 7%. The important factor influencing on specific surface area of phase interface is deformation of drops (bubbles) surface that in general case is caused by dynamic head under the effect of turbulent pulsations of disperse medium rate and (or) phases movement rate because of the difference in their densities. In this case the minimal size of dispersion phase particles dc, undergoing to deformation may be calculated from the ratio characterizing stability of phase interface (1.23) and (1.24). The dependence of minimal value of average diameter of disperse phase particles that subject to deformation at the expense of hydrodynamic effect of continuum medium was received by numerical experiment data analysis: 0 0,6 -0,4 (3.43) dc,=0,099*() *E p1 Diameter d,, may be accepted as diameter of forming diaperse particles in turbulent flow d2. For verification of adequacy of suggested method for disperse phase particles sizes calculation experimental investigation of emulsification process in turbulent flows limited by impenetrable wall of divergent-convergent design in hexane-water (continuous phase) system was carried out (see 2.2.7). Six-sectional tubular apparatus differ in canal geometry were used (Table 2.1). Experimental data on hexane emulsification effectiveness in reaction zone of divergentconvergent design confirm calculation results received with the help of the worked out technique (Fig. 3.31) very well. Equations (3.41) and (3.43) are held for values of densities of both continuous and disperse phases in interval 0,68911111 kg/m3. Thus, under studying of multi-phase flows in reactor of divergent-convergent design the results reliable enough can be obtained by the method of interpenetrative continuums (Eiler's approach). Numerical solution of differential equations of k-E turbulence model in partial derivatives with the help of implicit integral-interpolar method of final volume allows obtaining along apparatus volume of divergent-convergent design the fields of following functions: axial u and radial v rates for each of phases, pressure p, volumetric fractions of continuous a1 and disperse a2 phases, turbulence kinetic energy density K, its dissipation E and of some other characteristics.

79 d2. mm

d2,mm 19

1.4

os 1

OB

0.5

80

100

120 w1,um3(@5

1.30

181

b) a.mm

62, mm

2.4

1.4

16

1

0.8 40

io

w1, sm3lsec

100

0.6

C)

Figure 3.31. The dependence of averaged by volume of reaction zone diameter of drops dz on disperse medium movement rate w1.Points - experiment, line - calculation. Apparatus number (see Table 2.1): 1 (a), 2 (b), 3 (c), 4 (d). The method of calculation of turbulent mixing under multi-phase flows in reactor of divergent-convergent design allowing calculation the quality of resulted disperse systems "liquidliquid" and "gas-liquid" under fast chemical and mass-transfer processes proceeding was suggested. Analytical expression for calculation of average diameter of disperse phase particles forming in apparatus of divergent-convergent design was received and its adequacy was confirmed experimentally. For perfection of concrete technological studies when the process proceeds at the conditions of phase interface presence (polymerization process stop, polymer cleaning from catalyst leavings) analytical dependences of quality of disperse systems on geometrical parameters of mixing zone are very useful. As a consequence experimental studying of regularities of emulsions formation from two immiscible liquid flows in tubular turbulent apparatus of divergent-convergent design differing in canal geometry is very important.

80

3.2.2. The apparatus geometry effect As much as for compact tubular turbulent apparatus of cylindrical design (dd / d, = 1, L, / dd = 40) in canals of divergent-convergent design under two immiscible liquid flows movement rate w distribution of disperse phase drops by sizes converges and curves are shifted to the region of homogeneous thin-disperse systems formation. On the base of experimental data the equations correlating emulsification effectiveness (percentage content n of disperse phase drops with diameter d IO,8mm) in tubular canals of divergent-convergent design with flows rate w (R - correlation coefficient, apparatus numbers are in Table 2.1) were obtained: n = 0,78.w-51,0 (apparatus No. 1, R=0,96) (apparatus No.2, R=0,98) n = 0,88.w-58,2 n = 0,90.w-30,1 (apparatus No.3, R=0,99) (3.44) n = 0,72.w-35,6 (apparatus No.4, R=0,97) The rise of flows movement rate w and divergent-convergent section number N, from 1 to 4 leads to the reduction of surface-volumetric diameter of disperse phase drops and correspondingly to the increase of specific phase interface (Fig. 3.32) that in the case of fast processes intensifies proceeding of chemical process in the whole. Irrationality of application of tubular turbulent apparatus with divergent-convergent sections number N, more than four makes these devices compact (their length doesn’t exceed 8-10 calibers: L/dd) and simple in manufacture and exploitation.

Figure 3.32. The dependence of d32 on rate of heterophase flow and divergent-convergent sections number N, for apparatus No.1 (see Table 2.1). The dependence of d32 on rate of heterophase flow in divergent-convergent canals of various structures is presented in Figure 3.33. The equation approximating this dependence was received under statistical processing of experimental results (dd = 24 mm): d32 = (4,74).e-‘0’011’’w (3.45) It is obvious from comparison of experimental and calculation data (Fig. 3.33) that the ratio (3.45) describes the influence of w on disperse phase particles sizes d32 with a fair degree of accuracy (average discrepancy is 6%). At the same time there is a definite interval in volumetric rate of hetero-phase flows to which the divergent-convergent canal with optimal ratio dd / d, is corresponded. The zone of fiber movement of hetero-phase flows limits this interval from below,

81

and energy consumptions appearing as a consequence of pressure fall at apparatus ends in accordance with ratio Ap - V2 limits the interval from above. In particular, the interval 44 < w < 80 cm3/sec corresponds to ratio dd / d, = 3, and interval 80 < w < 180 cm3/sec to dd / d, = 1,6. The further rise of disperse system movement rate (w > 180 cm3/sec) determines obviously the necessity of further reduction of dd / d, ratio (Fig. 3.33) down to dd / d, = 1, i.e. compact tubular apparatus of cylindrical design in this case are effective enough. The same limitations determine boundary condition of dependences (3.44) application. According with equation (1.21) characteristic change of specific phase interface (Fig. 3.33) determining the mass-exchange process intensity is observed. d32, mm

20

F 10-4. mm-1

40

60

80

100

120

140

160 w, i&sec200

Figure 3.33. The dependence of surface-volumetric diameter of disperse phase particles d32 and specific phase contact surface F on geometry of apparatus of divergent-convergent design and reaction mixture movement rate. Apparatus (see Table 2.1): 1 (+V), 2 (mo), 3 (AA), 4 (00) experiment; lines - calculation by equations (1.21) (F), (3.45) (d32). Thus, in tubular turbulent apparatus of divergent-convergent design in comparison with cylindrical canals the flow in which disperse phase particles are uniformly distributed in reaction zone volume is formed at lower disperse phase movement rates, and the higher the dd / d, ratio the lower the required value of w . Polydispersity coefficient k D (1.22) is one more important magnitude characterizing resulted emulsions. For monodisperse systems - k D = 1 and for polydisperse k D < 1, and the lesser the k D parameter value, the bigger the dispersal of particles by their sizes. The ratio L, / & practically doesn't influence on resulted emulsions polydispersity as in the case of dispersions formation with surface-volumetric diameter d32 (Fig. 3.33). Particles dispersal is increased when dd / d, ratio is raised and sufficiently homogeneous emulsions are formed in divergent-convergent canal of tubular apparatus with dd / d, = 1,6. In particular for Ls/ dd = 2-3 the value of k D at dd / d, = 1,6 is 0,720,75, where as at dd / d, = 2 and 3 k D is decreased down to 0,63 and 0,41 accordingly. In general case in relation to influence of canal's geometry in tubular turbulent apparatus on regularities of disperse systems we may come to the following conclusions: Increase of rate of disperse system movement and number of divergent-convergent sections from 1 to 4 leads to the rise of specific phase interface in tubular turbulent apparatus that allows using of agitators with length equal to 8-10 calibers; Increase of dd / d, ratio in tubular turbulent apparatus of divergent-convergent design determines reduction of disperse system movement rate required for formation of flow with

82

uniform disperse phase distribution. By this, in the case of comparable apparatus productivity under others equal conditions polydispersity of formed emulsions is increased; In interval of / & ratio from 2 to 3 in tubular turbulent apparatus emulsions with comparable specific phase contact surface F and polydispersity k D are formed. Thus, application of compact tubular turbulent apparatus of divergent-convergent design allows effective dispersion in systems "liquid-liquid", real increase of specific phase contact surface and consequently reduction of diffusion limitations under fact processes proceeding under polymers synthesis, in particular under decomposition and cleaning out of catalyst, etc.. Change of liquid flows rate V in tubular apparatus and dd / d, ratio is practically the only but very effective method of influence on dispersion character and resulted heterophase reaction systems quality. This allows easily to control both fast chemical and physical mass-exchange processes proceeding in emulsions. The obtained regularities without technical and technological problems allow creation of thin homogeneous dispersions in systems "liquid-liquid'' if necessary. They also allow application of simple compact tubular turbulent apparatus of divergent-convergent design for intensification of chemical and mass-exchange physical processes first of all at the expense of thin-disperse systems formation under optimal conditions at minimum reagents residence time in mixing zone. We may recommend to use tubular turbulent apparatus of divergent-convergent design in industrial conditions both as main reactors for fast chemical heterophase reactions carrying out (chlorination of butyl rubber in solution) and as prereactor for formation of thin uniform emulsions with their following introduction into stirred tank reactors (SKEP(T) production). This technique will obviously allow to decrease time necessary for heterophase processes proceeding and to increase the resulted polymer products uniformity. In spite of canal geometry in tubular turbulent apparatus the way of reagents flows introduction significantly influence on quality of disperse systems. As a consequence, it is necessary to consider the influence of way of immiscible liquid flows introduction into tubular turbulent apparatus of divergent-convergent design on effectiveness of disperse systems formation process in turbulent regime.

3.2.3. The reagents introduction method effect The dependence of surface-volumetric diameter of disperse phase drops d32 on volumetric rate of disperse medium w has exponential form (Fig. 3.33) and with correlation high enough is straighted in logarithmic coordinates ln(d32) = f(w) (Fig. 3.34). Numerical dependences on the influence of volumetric rate of disperse medium w on size of disperse inclusions in particular surface-volumetric diameter were received on this base in dependence on flows introduction method (apparatus numbers are in Table 2.1): In(d32) = -0,0121 w +1,161 (apparatus No.5, R = 0,96) (apparatus No.6, R = 0,98) In(da) = -0,0129 w +1,119 (3.46) (apparatus No.7, R = 0,99) In(d32) = -0,0114 w +0,816 (apparatus No.8, R = 0,98) In(d32) = -0,0216 w +0,145 It is obvious from Figure 3.34 that increase of dl /d2 ratio at dl = const (apparatus NOS-8) leads to decrease of surface-volumetric diameter of drops and accordingly to (1.2 1) to the increase of phase contact surface. It possible occurs due to the fact that final size of disperse inclusions is determined by initial drops size. The last one in its turn significantly depends on diameter of floating sucker dz. Variation of ratio dl /dz at dz = const (apparatus No.7 and 8, Table 2.1) significantly influences on disperse structure if dispersions in tubular turbulent apparatus in particular reduction of dl /d2 (decrease of dl) leads to the formation of systems with minimum disperse phase drops sizes at given experimental conditions (Fig. 3.34). Significant decrease of disperse inclusions sizes in this case at comparable process productivity is determined by increase of disperse medium movement linear rate in sucker dl in 4 times that increases the intensity of shifting deformations to the drops of disperse phase at it radial way of introduction (Fig. 2.4). Moreover, the rate of drops

83

sizes change while changing of disperse medium consumption (flow rate) in accordance with (3.46) is in 1,8-1,9 times higher for apparatus No.8 than for apparatus No.5-7. ha321 h m l

-1

-

-2

-

-3 J

Figure 3.34. The dependence of h(d3z) on w. Apparatus No.5 (l),6 (2), 7 (3), 8 (4)(see Table 2.1). It was revealed that in dependence on liquid flows introduction method local hydrodynamics resistances of divergent-convergent design work differently. In particular, for tubular turbulent apparatus No.7 and 8 due to formation of fine-dyspersated emulsions in places of initial flows introduction the size of disperse inclusions practically doesn't change along the length of mixing zone (Fig. 3.35). In this case geometry of tubular turbulent apparatus of divergent-convergent design providing the high level of turbulent mixing as moving off introduction point of initial reagents (Fig. 3.6b) maintains the size of disperse inclusions on constant level at the expense of prevention of their coalescence and disperse system aliquation (Fig. 3.35, curves 3, 4). For apparatus No.5 and 6 (Table 2.1) due to comparably large diameters dz in points of flows introduction into apparatus big drops are formed that smashing as moving from the 1" to the 4'" divergent-convergent sections (Fig. 3.35, curves 1, 2 ) . In this case divergent-convergent sections along with providing of uniform disperse flow also operate as the dispersing devices only at the expense of geometry of canal itself. Application of apparatus of cylindrical design in such cases is unsuitable. Increase of homogeneity of received disperse systems with increase of volumetric discharge of disperse medium w (Fig. 3.36) is characterized for considered tubular turbulent apparatus of divergent-convergent design (Table 2.1). Apparatus No.5 with high value of diameter of feeding sucker d2 = 3mm is characterized by obtaining of comparatively uniform disperse system with k D = 0,86+0,94 in the whole considered interval of disperse medium rates. The value of dispersion of emulsion particles obtained in tubular turbulent apparatus No.6 and 7 is practically the same as rate of disperse system introduction with k = 0,63 at w = 44 cm3/sec up to k = 0,91 at w = 132 cm3/sec (Fig. 3.36) is increased.

84 d32,mm

1

0

t 1

1 !5

25

2

3

3,5

4

4 s

pJ

5

Figure 3.35. Dependence of surface-volumetric diameter d32 on sections number. Apparatus No.5 (l),6 ( 2 ) , 7 (3), 8 (4). w = 132 cm3/sec.

40

EO

120

160

w , sn 3bec

Figure 3.36. The dependence of polydispersity coefficient kD on w. Apparatus No.5 (I), 6 (2), 7 (3) (see Table 2.1).

85

Thus, the influence of method of initial reagents introduction into tubular turbuleni apparatus on quality of resulted disperse systems "liquid-liquid'' was revealed. Numerical dependences on the influence of construction of feeding suckers and apparatus productivity on surface-volumetric diameter of disperse phase drops were received. The influence of dl I d2 ratio variation in the case of dl = const or dz = const on disperse phase drops size is slope opposite. In particular the increase of dl Id2 ratio at dl = const and its reduction at dz = const lead to obtaining of fine-disperse emulsions with developed surface of phase contact. At dl I dz < 10 (dl = 1Omm) divergentconvergent sections in tubular turbulent apparatus operate as dispersing devices and reduce possibility of coalescence processes proceeding. Under realization of processes at phase interface in conditions of industrial production it is advisable to use tubular turbulent apparatus of divergentconvergent design with marginally possible diameters of sucker for introduction of disperse medium and radial placed sucker of disperse phase introduction. It is well known that multiphase flow in tubular canals is accompanied by fibering. Phase disengagement leads to reduction of specific surface of reacted flows and consequently to additional diffusion limitations under fast chemical and mass-exchange physical process realization under polymers synthesis. This determines the expediency of determination of boundary condition of homogeneous and fibered regions of reaction mixture flow formation in tubular turbulent apparatus. 3.2.4. Separation of reaction medium under multi-phase systemsjlow in tubular turbulent apparatus At low rates of disperse system movement in tubular turbulent apparatus of both cylindrical and divergent-convergent designs fibered flows are formed (Fig. 3.37a). The rise of disperse system rate leads to the formation of homogeneous flow with uniform distribution of disperse phase particles along the whole reaction zone volume (Fig. 3.37b). In tubular turbulent apparatus of cylindrical design in practically the whole considered interval of rates of dispersion medium the formation of fibered flow is observed (Fig. 3.38). However, this region (111, Fig. 3.38) converges as w is increased. In the case of apparatus of divergent-convergent design fibered flow of reaction mixture is formed in sufficiently narrow region and only at low rates of flows (w < 50 cm3/sec) (112, Fig. 3.38). Inclination to fibering is determined by initial diameter of disperse phase drops. In particular under increase of side sucker diameter d2 for disperse phase introduction (Fig. 2.4) the rise of drops sizes is observed (Fig. 3.35) and reduction of homogeneous flow region extension occurs (Fig. 3.39).

Figure 3.37. Formation of fibered (a) and homogeneous (b) flows in system "water-hexane" in tubular turbulent apparatus.

m

I

0

20

40

60

80

100

120

Figure 3.38. The regions of homogeneous (I) and fibered (11) flows of reaction mixture in apparatus of cylindrical (1, apparatus No.9, see Table 2.1) and divergent-convergent (2, apparatus No.5, see Table 2.1) designs.

Figure 3.39. Dependence of transition limit "homogeneous-fibered flow" in reaction zone of divergent-convergent design on diameter of side sucker dz (apparatus No.5-7, see Table 2.1).

87

Phases separation in tubular turbulent apparatus of divergent-convergent design occurs also at high rates of reaction mixture movement that determine the superior limit of process productivity. As it was mentioned above near the impenetrable divergent-convergent wall under effect of pressure gradient flow's particles begin to move opposite to main flow (Fig. 3.29). In this region liquid flow moves rotationally that leads to the appearance of unusual effect: under the centrifugal forces action the separation (division) of phases occurs and it firstly depends on density of reagents participating in dispersing process. Under significant differences in densities of continuous p1 and disperse p2 phases and p1< p2 (for example system "liquid-solid") in peripheral apparatus part reduction of volumetric fraction of disperse phase a2 is observed at the expense of particles dropping to the impenetrable wall or to periphery under centrifugal forces action (Fig. 3.40). In the case of p1> p2 (for example system "liquid-gas") decrease of volumetric fraction of disperse phase az in circulation zone is observed. Under insignificant difference in densities of dispersion medium and dispersion phase ( P I = p2) separating effect is not observed and disperse phase is uniformly distributed along reaction zone (Fig. 3.37b). Homogeneous two-phase flow formats also in narrow section of apparatus (convergent). In this case there is regular distribution of volumetric fraction of disperse phase a2 along apparatus section (Fig. 3.40). Carried out calculations allow making the following conclusion: under the use of tubular turbulent apparatus of divergent-convergent design it is necessary to consider the possibility of phase separation effect appearance and formation of heterogeneities in disperse system components distribution under realization of both chemical and mass-exchange physical processes. a2

2

0.15

n

n6

0

0,02

r,

0.04

Figure 3.40. Distribution of volumetric fraction of disperse phase a2 by apparatus radius in divergent (1, 2) and convergent (3, 4) parts. Disperse medium density p1= 6,93 (1,3); 1111 (2); 998,2 (4) kg/m3. Disperse phase density p2 = 640 (1,3); 0,689 (2); 1111 (4) kg/m3. (dd = 0,08 m, d, = 0,05 m, Ls= 0,27 m, V = 16 d s e c ) .

Centrifugal acceleration of particle got into swirling flow is estimated as a = Vz / R and rotation radius is assumed to be equal to R = (dd - dc) / 2. Presenting Archimed's criterion in modified form A r d and substituting the gravitation sexpressed as acceleration of gravity g in it by centrifugal acceleration we shall receive: 3 2 2 2d2P1 v (P, - P2 1 (3.47) = 2 P1 (dd - dc )P,

Armed

88

On the base of analysis of physical picture analytical formula was suggested for estimation of separation effect in which numerical coefficient and index of degree were identified by numerical calculation results processing: (3.48) where Act2 -maximum difference of volumetric fractions of disperse phase on apparatus axis and on periphery; azav- average value of volumetric fraction of disperse phase in reaction flow; Re = Vdzpl/ p1. Thus, in tubular turbulent apparatus of divergent-convergent design in comparison with cylindrical design fibered flow is observed in region narrow enough and only at low rates of flows movement. Formation of homogeneous two-phase flow in apparatus of cylindrical design is observed at high linear rates of reaction mixture movement (V > 0,7 d s e c ) . Possibility of initiation in peripheral part of tubular turbulent apparatus of divergent-convergent design of separation effect at high linear rates and significant difference in multi-phase flows reagents densities was revealed. Criterion of formation of separation effect in dependence of zone geometry of flows mixing and their physical parameters was obtained. Carried out investigations allow calculation of optimal operation working regime of turbulent apparatus with consideration of possibility of phases separation effect appearance and formation of heterogeneities in disperse phase components distribution. The last fact was realized under ethylene with propylene copolymerization on ZieglerNatta catalysts and butyl rubber chlorination by molecular chlorine. Fundamental regularities of turbulent mixing of one- and multi-phase flows worked out in this chapter allow significantly deepening and extending of novel part of theoretical technology concerning fast chemical processes proceeding in turbulent regime. The analysis of physical picture of reaction mixture turbulent movement in tubular canals of various construction determines the possibility of quantitative and scientifically-founded approach to the selection of optimal reaction zone geometry under any fast chemical processes proceeding with the aim of reduction of diffusion limitations and creation of uniform conditions under synthetic products production. However, search of solutions of one more large scientific and applied problem connected with temperature regime control under fast chemical processes realization with local evolution of significant amount of heat is necessary. There are still no clear methods of effective use of external thermostating of fast exothermal reactions.

89

CHAPTER 4. EXTERNAL HEAT REMOVAL UNDER FAST CHEMICAL REACTIONS IN TURBULENT REGIME

4.1. Quasi-plug-flowmode in turbulent flows as a key to realization of fast chemical processes in quasi-isothermalconditions The possibility of formation of quasi-plug-flow mode in tubular turbulent apparatus when reaction zone approaches heat-exchange surfaces determines possibility of thermostating of fast chemical reactions by external heat removal (see 1.2.1.4 and 1.3.3.3). At the same time there are no data about the influence of fast chemical reactions kinetic parameters and reacted flows physical characteristics (density, viscosity) on conditions of this practically important regime of tubular apparatus operation. Meanwhile there are often a lot of significant complications of fast processes carrying out in industrial production, for example under cationic piperylene oligomerization in the presence of A1C13.0(C&)z as a catalyst, under butyl rubber chlorination, neutralization, etc..

4.1.1. The influence of fast chemical reaction kinetic parameters on conditions of quasi-plug-flow mode formation in turbulent flows The possibility of formation of corresponding liquid flows mixing fronts under fast chemical reaction proceeding and without it (see 1.3.1) determines the interest in investigation of reaction rate constant effect on conditions of quasi-plug-flow mode formation in turbulent flows (plan reaction front). Low-molecular chemical reactions of the second order proceeding with rate constants in the range of k = 102i108l/mole.sec were experimentally studied. It was found that chemical process contributed significantly in conditions of corresponding macroscopic reaction fronts formation under mixing of reacting flows. This fact is revealed in reduction of requirements to macroscopic structures formation conditions: increase of chemical reaction rate is analogous to rise of turbulence level in reaction zone and improves reagents mixing process (Fig. 4.1). As a consequence, low bounds of characteristic macrostructures are fixed at lower values VI / V2. For low bound of quasi-plug-flow mode formation under proceeding of fast chemical reactions of the second order the following ratio is right (at dl / dd = 0,44): Vi / Vz = -00'ilgk + L6 (2 < lgk < 8) (4.1)

=JD6 /IdC]"-' shows that under increase of chemical CB reaction rate constant k in particular under fast chemical reaction of low-molecular compounds synthesis and also under rise of reagents concentrations the values of Kr reduce. Reduction of &r as criterion of quasi-plug-flow mode formation in turbulent flows allows decreasing of Dt value determined by linear rate of reacted flows, in particular by ratio V I / V2. The rise of reaction order can change requirements to formation of plant front towards one or another side: Rcr reduces at [C] > 1, and increases at [C] < 1. Fast chemical reaction proceeding also influences on the conditions of formation of low bound of torch front (Fig. 4.1, curves 5, 6). Under rise of value of chemical reaction rate constant k the ratio of linear rates of reagents introduction into apparatus required for formation of torch regime low bound reduces. One may suggest that kinetic parameters of chemical reaction (rate constant in this case) don't change the region of corresponding macro-structures formation. Only the displacement of ratio V I / V2 of rates required for formation of quasi-plug-flow and torch regimes to the region of lower values occurs. Analysis of expression (1.13) R

90 VlNZ

VZ,m$ec

0

0,15

0,35

OS5

Figure 4.1. Conditions of quasi-plug-flow mode (1-4) and torch front ( 5 , 6) formation under proceeding of following reactions: neutralization (1); interaction of potassium thiocyanate with iron chloride in acid ( 2 ) and neutral (3, 5 ) mediums; mixing of methyl green solution with water (4,6). (dl / dd = 0,44, see 2.2.5) Thus, it is obvious that under mixing of liquid reagents at conditions of chemical process proceeding kinetic and diffusion factors are of great importance and this fact is expressed in the influence of rate constant of proceeding reaction on formation conditions of topochemical regimes in tubular turbulent apparatus. It is interesting to note that under calculation of turbulent flows with fast chemical processes the influence of chemical reaction rate constant on effective diffusion coefficient and micro-mixing accelerating at the expense of local reagents concentrations gradients increase was predicted [ 1371. The dependence of low bounds of reaction fronts characteristic macro-structures formation, in particular of plant front and torch on the values of proceeding chemical reactions rates constants is the experimental evidence of correlation between diffusion coefficient and chemical reaction rate constant. At the same time we may suggest that formation of characteristic macro-structures of reaction fronts is determined by mixing both at macro- and microlevels. As in industry including butyl rubber chlorination and piperylene cationic oligomerization we often have to use liquid flows with different physical characteristics (density, viscosity) it is necessary to study the influence of these parameters on conditions of reaction plan front formation - quasi-plug-flow mode, and consequently on the effectiveness of external heat removal. 4.1.2. The influence of liquidjbwsphysicd characteristicson conditions of quasi-plug-flow mode formation

Change of density and viscosity of liquid flows introducing into tubular turbulent apparatus significantly influences on conditions of quasi-plug-flow mode formation and consequently on effectiveness of external heat removal under fast polymerization processes in turbulent regime. Rise of density of even one reagent or solvent leads to the reduce of the required ratios of linear rates Vi I Vz necessary for the formation of quasi-plug-flow mode (Fig. 4.2). Rise of viscosity of liquid flows is accompanied by increase of V1 / VZ ratio. In other words, for effective mixing and thermostating in tubular turbulent apparatus and for its optimal functioning reagents with different viscosities should be introduced with higher rate of axial flow Vi (Fig. 4.3.). The character of influence of density and

91

viscosity of mixing (reacted) flows on formation of plan reaction front determining process stationarity and external heat removal effectiveness in turbulent flows limited by impenetrable wall is determined in accordance with data that presented in Figures 4.4 and 4.5 by following dependences: V1 I Vz = 4 9 p + 8,2 (p = l+1,2 g/cm3, di / dd = 0,44) (4.2) Vi I V2 = 1,07p + 6,2 (p= l+6 mhSeC, dl / dd = 0,13) (4.3) JllV2

0.1

0,3

0.5

V2. mloec

Figure 4.2. Conditions of torch (5) and quasi-plug-flow (1-4) regimes formation in dependence on flows densities: 1,02 (1); 1,07 (2); 1,12 (3); 1,19 (4); 1,02, 1,07, 1,12, 1,19 (5) g/cm3. (dl / dd = 0,44, see 2.2.5).

04 0.35

0.5

0,65

w. w5RC

Figure 4.3. Conditions of torch (5) and quasi-plug-flow (1-4) regimes formation in dependence on flows densities: 1,3 (1); 2,5 (2); 3,8 (3); 6 (4); 1,3, 2 5 , 6 (5) mPa.sec. (dl / dd = 0,13, see 2.2.5).

92 V1N2

1

1,05

1,l

1,15

r, glsm3

1 2

Figure 4.4. The regions of torch ( 2 ) and quasi-plug-flow (1) regimes formation in dependence on flows rates and densities. VZ= 0,5 mlsec, dl I dd = 0,13. V1N2

v

.

0

2

4

6

m, mPa sec

Figure 4.5. The regions of torch ( 2 ) and quasi-plug-flow (1) regimes formation in dependence on flows rates and viscosities. VZ= 0,5 d s e c , dl I dd = 0,13.

93

It was revealed that the fact that liquid flows' densities and viscosities change doesn't influence on reaction torch front formation is also very important (Fig. 4.2, 4.3). The rise of density of flows at the expense of reduction of low bound of ratios V I / V2 required for the formation of plan front and constancy of ratios VI / VZrequired for formation of low bound of torch front leads to the narrowing of ineffective torch regime region formation. The last fact widens possibilities of tubular turbulent apparatus exploitation in optimal quasi-plug-flow regime (quasi-isothermal regime) (Fig. 4.4). Thus, the rise of liquid flows density leads to more stable work of tubular turbulent apparatus in high-effective quasi-plug-flow mode. And the flows viscosities rise leads down to unstable functioning of tubular turbulent apparatus of cylindrical design at the expense of increase of low bound of ratios V1 I V2 required for formation of quasi-plug-flow mode and constancy of ratios VI / V2 required for torch front formation (Fig. 4.5). The observed picture may be explained by the influence of density and viscosity of liquid flows on turbulence level determined in the first approximation by criterion Re (see 1.3.4). In particular, increase of moving reagents flows density leads to the rise of Re values, Le. hydrodynamic similarity of system is broken. To form in the apparatus quasi-plug-flow mode at changed reagents density it is necessary to obtain previous values of Re. This can be realized by decreasing of linear rate V of liquid movement or by reduction of apparatus diameter, and also by increase of system viscosity. In our case the hydrodynamic similarity of system and formation of quasi-plug-flow mode are reached by decreasing of ratio VI 1 V2 (reduction of axial flow rate). Effect of viscosity on plan front formation conditions may be explained analogously but with opposite conclusions. Application of tubular turbulent apparatus of cylindrical design with various diameters ratio d1 I dd showed that decrease of dl / dd lead to the fact that quasi-plug-flow mode was formed at lower volume flow of reagent introducing through axial sucker with diameter dl. In particular under reduction of dl / dd from 0,44 down to 0,13 the volume throughput of axial flow at which quasiplug-flow mode formation begins reduces by 60%. This fact allows application of more concentrated reagents solutions in chemical process that is very useful under neutralization of acid and alkali manufacturing waters. Thus, correlations (1.13) and (4.1-4.3) at homogeneous mixing of liquid flows including those differing in density and viscosity allow effective usage of tubular turbulent apparatus in industrial production practically at any stage limited by mass-exchange. They also allow calculating of conditions of stable operation of tubular turbulent apparatus in quasi-plug-flow mode with their optimization including optimization by change of physical characteristics of liquid flows, kinetic parameters of fast chemical reaction and also of apparatus construction. Obtained on the base of model low-molecular reactions regularities are also suitable for calculation of technological parameters under fast processes of high-molecular compounds synthesis, in particular piperylene oligomerization, butyl rubber chlorination in solution, etc. at turning on of external heat removal. Moreover the effect of chemical reaction rate constant and also of some physical parameters of liquid flows (density, viscosity) on conditions of characteristic macroscopic fronts formation in turbulent flows limited by impenetrable wall allows supposing the various nature of reaction and mixing fronts formation. In the first case kinetic and diffusion process parameters are determinant, and in the second - preliminary convective and turbulent transfer. The influence of density and viscosity of liquid flows, Le. parameters determining hydrodynamic regime of liquid flows in tubular canals on conditions of reaction and mixing plan front formation shows the important role of hydrodynamic constituent also in general case under corresponding macrostructures formation. Formation of quasi-plug-flow mode in tubular turbulent apparatus and consequently of optimal conditions for application of external heat removal requires creation in apparatus volume of high turbulent mixing level determining by value of Dt, and Dt significantly depends on reaction zone geometry (see 3.1.2). Thus, after investigation of dependence of turbulence level on apparatus construction and systems hydrodynamic parameters we have come to the conclusion that there was really one more

94

factor that allowed realization of fast chemical and mass-exchange physical processes under polymer synthesis at optimal plug-flow regime in turbulent flows. Guaranteed formation in reaction zone of quasi-plug-flow mode determines effectiveness of thermostating of fast polymerization processes at the expense of external heat removal. This determined the expediency of working out of the ways of temperature field control in reaction zone at turning on of heat removal through heat exchanging reaction zone surface.

4.2. Regulation of thermal conditions under fast chemical reactions in turbulent regime Maintenance of required temperature regime in reaction zone is a necessary condition of effective and safe carrying out of any process. Rise of temperature in reaction zone under proceeding of exothermal processes depends on thermodynamic and kinetic parameters, reagents concentrations and way of their introduction, hydrodynamic regime in reaction zone, etc. Problems in temperature regime control in reactors are aroused under carrying out of fast chemical exothermal reactions in liquid phase, which proceed for seconds and reaction zone length Lch by this doesn't exceed several centimeters and sometimes parts of centimeters (Table 4.1). Naturally it leads to local heating of medium and may lead to heat explosion if there is no heat removal. Realization of fast chemical processes in standard stirred tank reactors with any mechanical mixers and system of external heat removal is technologically inexpediently due to this reason because in reagents in-let places volume temperature gradient is always formed - "torch". It is characterized by non-stationary reaction proceeding, significant and very often inadmissible overheating of reaction mixture in reaction zone (see 1.1.4). Fast chemical processes should be carried out in flow, in compact tubular apparatus of jet type functioning in ideal plug-flow regime in turbulent flows when in interval of definitely limited numerical values of reaction zone radius plan front of reagents concentrations and temperature is formed, i.e. quasi-isothermal (isothermal in any reaction section) regime of tubular turbulent apparatus operation is formed (see 4.1). As a consequence it is necessary to study heat regime in reaction zone and to reveal the ways of effective regulating of temperature profile in tubular turbulent apparatus under realization of fast chemical processes at the expense of external heat removal in technically admissible conditions.

Table 4.1. Proceeding of fast exothermal processes in flow with the use of tubular turbulent apparatus of cylindrical design at condition of independence of heat effect. heat cauacitv and medium densitv on ternuerature iTn = 293 K. = 243 K, V = 1 d s e c ) No 1. 2.

6.

Process

I

ec

CH2=CH2+HCl+ CH2Cl-CH3 CH2=CH2+C12+ CH2CI-CH2Cl

H2S04+nH20+ H2S04.nH20 (n=3,6)

A92 1,4

107+~

0,130

t

I

1254

I

552

1238

0,147

1615

1899

1,335

0,71 1000

0,128

1725

1163

3,680

2,Ol 1200

0,165

5120

1970

3773

1,700

0,52

0,134

1637

1064

340

2,280

5,52 1498

0,165

6343

196

1,674

0,22

700

0,134

1014

97 1

1,575

I0,25

860

0,131

1114

2706

1,748

1113 1206

860

I

0,141

Notes: * 9***kmole/m3;** well-type cylindrical apparatus (S = 0,126 m2, R = 0,2 m); Shell-and-tube cylindrical apparatus (R= 0,Ol m).

~

,

I

1520

I

584

1

131

I

1 I 247

2.10-"

I I 957

5-10-2 I422,5

I

26,3 11,6

96

4.2.1. Selection of thermal conditions control mode under fast liquid-phase reactions in turbulent

regime on the account of external heat removal In dependence on numerical values of heat effect q of process and reaction product yield A l l temperature rise ATad in reaction zone according with (1.4) may be equal to tens and even hundreds of degrees. By this all heat is evaluated very quickly (for seconds or parts of seconds) and for a very little distance along reactor length (Table 4.1). As a consequence effective heat removal directly in reaction zone at the expense of external cooling is practically impossible, and one should always take this fact into account. For tubular turbulent apparatus of cylindrical design it is necessary to obtain the equation for calculation of temperature field in reaction volume under fast chemical reaction proceeding (supposing that reaction proceeds for 100% at places of reagents introduction). In this case heat balance equation for element of surface dF is as following: dQ = GICrr(-dTr) = GcCrcdTc (4.4) Here Gr, G,, Crr, Crc- mass flow and heat capacity of reaction mixture and coolant, accordingly. We receive the following expression out of (4.4) after corresponding transformations:

d(T - T )=-dQ(-+r c

1

(4.5)

At the same time according with main heat-transfer law dQ = KdF(Tr-Tc) equation (4.5) may be transformed to: 1 1 d(Tr - Tc) = -KdF(+) GrCrr GcCrc Then after we have divided variables, integrated the obtained expression within the limits of change from Tad - T,1 (Tad = To + ATad) to Tr - T,z (Tr - required temperature in reaction zone) and dF within the limits from 0 to F (at K = const) the following equation was obtained: Tr - T 1 1 In c2 =-W((4.7) Tad - Tcl Grcrr GcCrc where F = 27rRL,,1 and G = VnR’p. For calculation of required zone of cooling in tubular turbulent apparatus according with (4.7) it is necessary to know heat-transfer coefficient K that is calculated out of following equation [3]: 1 E= (4.8) 1 6 1 ’

+-

) 7

-+-+a

h a

1 2 where a1,a2 - heat-transfer coefficient of cooling and heating flow (coolant), correspondingly; 6 wall thickness (0,001m); h - heat conductivity coefficient (for fused silica 1,389 J/(m.sec.degree)). Heat-transfer coefficient is usually found in reference books (for water at 323 K a = 1801,44 J/(m2.sec.degree)). Dependences of length of cooling zone on annular flow rate wa that is experimentally received (curve 1, Fig. 4.6) and calculated by equation (4.7) (curve 2, Fig. 4.6) are presented in Fig. 4.6. It is obvious that at low rates of annular flow (of coolant) significant discrepancy between experimental and calculated values of LCwl is observed. At the same time at high wa this discrepancy is reduced and at coolant movement rate about 350 cm3/sec calculated Lcoolcoincides with experimental. Heat transfer coefficient will be totally determined by coefficient of heat emission of cooling internal flow a1 due to the fact that when wa is increased the value of a;? is risen (because turbulence level of annular flow is increased). Thus, the equation (4.7) can be transformed with consideration of following conditions:

91

1) coolant rate is high (G, + m), so we may accept that coolant temperature is T, = const (this condition is practically always met because tubular turbulent apparatus are compact); 2 ) heat transfer coefficient may be substituted by heat emission coefficient of reaction mixture under fast chemical reaction proceeding (in accordance with (4.8) heat transfer coefficient is determined by lower heat emission coefficient). In this case, after transformations of (1.14) with consideration of suggested assumptions we shall obtain:

LCOOL rn

0,95

-

. . .

,

8

8

8

8

2 4

8

+

OB5

0,75

-

7

8

i/ d

Figure 4.6. Experimental (1) and calculated ( 2 ) lengths of cooling zone in reaction zone of cylindrical design. Tad = 5OoC, T, = 4-5OC, a1 = 1801,44 J/(m*.sec.K), 6 = 0,001 m, h = 1,389 J/(msec.K), wa= 47 cm3/sec,R = 0,008 m, Cr = 4179,91 J/(kg.K), p = 1000 kg/m3. The dependence of experimental and calculated out of (4.9) temperature difference in reaction zone on flows rate is presented in Figure 4.7. It is obvious that with assumption that T, = const the equation (4.9) for sure describes experimental data. This fact determines the possibility of application of this equation for analysis and search of ways of effective temperature regime control under fast chemical reaction proceeding. The usage in (4.9) of heat emission coefficient of reaction mixture instead of heat transfer coefficient (1.14) allows significant simplification of temperature field calculation in reaction zone because heat emission coefficient can be found with the use of similarity theory.

98 DTcool,OC

04 0

100

200

wB,cm3lsec

Figure 4.7. Comparison of experimental (points) and calculated by (4.9) (lines) temperature difference ATcoOlin reaction zone of cylindrical design in dependence on reaction mixture rate WB. wa = 130 cm3/sec, L = 1 m, Tad = 5OoC, T, = 4-5OC, C X ~= 800 J/(m2.sec.K), a1 - reference data (nomograms). We may estimate the length of cooling zone at external heat removal required for maintenance in reaction volume of necessary temperature Tr by solution of (4.9) in relation to Lcool:

(4.10) The length LCoolis increased with the rise of Cr, p, R and V, and heat emission coefficient a also depends on these magnitudes. The value of in equation (4.10) may be expressed by criterion Nu = 0,023 (Re)',' (Pr)OS3(turbulent regime) with the use of corresponding equations of criteria Re (1.32) and Pr (1.37) (see 1.3.4). (4.1 1) After transformation of (4.10) taking into account (4.11) we shall obtain:

(4.12)

99

Then, calculating values of Cr, p, p and h in first approximation independent on temperature we obtain C,0,7p0Lo'Jp0'2 l0,04 hop'= A = const. Then:

Evidently coefficient A is determined only by physical parameters of reacted system and practically doesn't depend on kinetic and hydrodynamic process parameters and for concrete chemical process there are all numerical values of these magnitudes significantly influence on the length of cooling zone L,,,l (Table 4.1). Equation (4.13) determines possible ways of temperature regime control in tubular apparatus of jet type under fast chemical processes at the expense of external heat removal, in particular by change of To, All, T,, R and V at condition of turbulence of reacted flow. Sufficiently simple method of regularization of length of cooling L,,1 is variation of reaction zone radius R and/or linear rate of reagents movement V however to significantly lesser degree. Accordingly to (4.13) increase of V from 1 up to 10 d s e c leads to the rise of LCmlin 1,6 times and change of R in order leads to the change of Lfwl in more than 15 times. Dependence of reaction medium temperature change (T,) in tubular apparatus on length of cooling zone (LC,1) on the example of fast reaction of ethylene hydrochlorination at various values of reaction zone radius R is presented in Figure 4.8. Reduction of temperature for 50°C in apparatus with R = 0,2m under reaction mixture cooling (T, = 283 K) is possible at L,,,,,l 2 229 m. When one uses coolant with T, = 243 K value of LCmlis decreased nearly in 2 times (L,,,,,l = 11Om) but remains very significant. However under use of tubular turbulent apparatus with less (for order) reaction zone radius R=0,02m the length of cooling zone Lc,,l is decreased down to 14,3 and 7 m correspondingly, i.e. in more than 15 times that is technologically enough. The obtained results unambiguously testify that under realization of fast exothermal processes (k 2 10'' l/mole.sec) in tubular turbulent apparatus of jet type in all cases removal of heat is practically possible only after processes are over (from reaction product) but not directly in reaction zone. It is very important because to reach isothermal regime directly in reaction zone when L, = Lcwlis practically impossible even at significant decrease of T, (Fig. 4.9). Thus, according with (4.13) decrease of reaction zone radius R is the effective way of heat regime control during tubular reactor functioning under fast chemical processes proceeding in turbulent flows (Fig. 4.8). However, for more radical improvement of parameters required for effective proceeding of fast chemical processes (Table 4.1) it is better to use shell-and-tube turbulent apparatus of jet type consisting of N pipe bundle with small diameters, washed by coolant (model of standard shell-and-tube heat exchanger) and operating in quasi-plug-flow mode in turbulent flows. In this case:

100

350

1

300

-

___ 0

5

10

20

15

25

30

Figure 4.8. Temperature change in reaction zone (Tr)along its length (L) under chlorinated ethyl synthesis at conditions of external heat removal. R = 0,2 m (1, 2); 0,063 m (3, 4);0,02 m (5, 6); 0,014 m ( 7 , 8); 0,Ol m (9, 10). T,= 283 K (1, 3, 5, 7 , 9); 243 K (2, 4, 6, 8, 10). ( V = 1 M/C, To = 273 K, A l l = 204 kg/m3, AT,d = 80').

400

-

300

-

1

2

3

200.

100

-

0

0

0.04

0,08

0.12

R,m

Figure 4.9. Dependence of cooling zone length ( Ll on) reaction volume radius R. Neutralization (An =204 kg/m3) (1, 2); dichlorethane synthesis (All = 67 kg/m3) (3, 4); chlorbenzene synthesis (An = 100 kg/m3) (5, 6); ethyl chlorine synthesis (AII = 204 kg/m3) (7, 8). Tc = 283 K (1, 3, 5, 7); Tc = 243 K (2,4,6,8). (V = 1 d s e c , TO= 273 K, ATad= 80°C)

101

Flow breaking under application of shell-and-tube turbulent apparatus of jet type to n flows with holding of tubes section S allows sharp reduction (in No" times) of cooling zone length Lc,,,,l (Table 4.1) that makes the process technological and technically simple. Values of Lo0l (by the example of ethylene hydrochlorination) and R for various number of tubes providing at given productivity of shell-and-tube reactor the temperature about 273 K in reaction zone at two values of T, (238K and 258K) are presented in Table 4.2. Table 4.2. Dependence of cooling zone length Lcool(T, = 273 K) on tubes number N and radius Rn in shell-and-tube reactor under ethylene hydrochlorination (conditions as in Table 4.1, V = 1 d s e c , To = 263 K, AT = 100°C)

It is obvious that maintenance of temperature at 273 K in reaction volume with R = 0,17m even at T, = 238 K is practically impossible because cooling zone length should be not more than 90 m. However when use shell-and-tube reactor maintenance of temperature at 273 K is possible from tubes number N 2 60 even at T, = 258 K (Lcool= 7 and 12 m, correspondingly). Moreover, potential possibility of process proceeding at conditions close to isothermal is appeared (L = 10 m, N = 100). Thus, in general case under fast liquid-phase chemical processes in tubular turbulent apparatus of jet type the reduction of reaction zone radius R and also its breaking to N constituents is the simple and effective way of heat regime control in reaction volume. Nevertheless very often it does not solve the problem of local overheating of reaction mixture at places of reagents introduction. So, to avoid inadmissible overheating of reaction mixture under fast chemical processes it is necessary to use "zone" model presenting consistent combination of several independent reaction zones. Reaction mixture comes to every reaction zone from previous zone after cooling down to T, and where new definitely dosed portion of reagent (catalyst) is introduced. This portion determines permissible temperature rise that in its turn is determined by reaction product yield in every zone. Tubular turbulent apparatus of jet type in this case presents an order of adiabatic mini-reactors of ideal plug-flow in turbulent flows with L, divided by zone of heat removal Lcool. Functioning of tubular turbulent apparatus at conditions of "zone" model is considered by example of interaction of sulphuric acid with water, in particular under dilution of 90% sulphuric acid down to 60% solution that is accompanied by evolution of large amount of heat (Fig. 4.10). High enough rate of interaction of sulphuric acid with water k = lo7*' I/mole.sec) and large reaction heat effect (q = 196 kJ/kg) determine adiabatic temperature rise for -84°C in very small reaction zone (L, = m) (Table 4.1). In this case in particular technologically acceptable conditions of process carrying out (temperature in reactor should not exceed 70°C (343 K)) may be obtained by application in given case of two-zone model of apparatus functioning with portion introduction of 90% acid solution into two zones divided by cooling zone. This approach (Fig. 4.10) allows process realization in technologically acceptable temperature regime. Moreover, portion acid

102

introduction determines possibility of reduction of apparatus radius in the first zone (from 0,Ol down to 0,008 m) that together with maintenance of required turbulence level of flows allows increasing effectiveness of external heat removal also (L'

is decreased for 1520%). Maximum cool permissible temperature rise of reacted system in tubular apparatus up to some critical value determines the number of necessary reaction zones under portion reagents introduction, i.e. necessity of "zone" model application. or, E

I

Izonc R=O.OOBm

!

360

320 2 1

I

2

1

L

L

COO1

coo1

280

0

5

10

15

20

L,m

Figure 4.10. Temperature change (Tr) along apparatus length (L) at conditions of external heat removal by the example of interaction of sulphuric acid with water in turbulent regime with the use of one- (1) and two-zone (2) models. (To = 283 K, Tc = 283 K, V = 1 m/sec, flow parameters are in Table 4.1). The better results under fast chemical reactions carrying out in liquid phase can be obtained at conditions of use of "zone" model under dividing of reaction zones by shell-and-tube apparatuscondensers. The model of work is considered by the example of cationic polymerization of isobutylene (Table 4.3). We should note that hydrodynamic regime of reaction mixture movement significantly influences on effectiveness of external heat removal in tubular apparatus. In particular, under transition from laminar to turbulent regime of reagents flows at condition of constancy of process productivity w (w = VmR2 = 10,3 m3/h was chosen) effectiveness of external heat removal is increased. As a consequence the length of cooling zone is significantly reduced (Fig. 4.1 1) and the increase of numerical values of Re from 2300 up to 4.104 at constant process productivity leads to the decrease of tubular reactor volume, in particular for water in 1000 times, for chlorethyl (under fast chemical reaction of ethylene hydrochlorination in turbulence regime) in 300 times. At unstable (transition) regime of tubular apparatus work necessary cooling zone length Leool is sharply increased even in comparison with laminar regime of liquid flows. One should always take into account mentioned facts under realization of fast chemical processes of polymers synthesis.

103

Table 4.3. Temperature regime and parameter of cooling zone under isobutylene polymerization in turbulent regime

$?polant) one-tube condenser with constant diameter (radius 0,2 m and section S, washing by coolant)

Lcool,

( ,Jb2sxE

m

2000 1

500

0

0

0

100

200

300 Re

Figure 4.11. Dependence of cooling zone length LCmland heat emission coefficient a on hydrodynamic regime of tubular apparatus work for water (0)(heat exchange) and chlorethyl ( 0 ) (heat exchange under proceeding of chemical reaction of liquid-phase ethylene hydrochlorination) at fixed productivity (10,34 m3/h). (Tad = 374 K; TO= 278 K; T, = 293 K; Te = 283 K). Moreover since under fast chemical processes (isobutylene cationic polymerization, acid medium neutralization, etc.) in tubular apparatus the length of reaction zone L, doesn't exceed

104

several centimeters and sometimes parts of centimeters (see Table 4.1) as it is mentioned above, so for realization of such processes one may recommend to use tubular apparatus with ratio L / R e 100. This fact determines rise of effectiveness of external heat-exchange (heat emission coefficient) in more than 1,5 times (up to 1,65 times) at the expense of additive turbulization coursed by disturbances of in-let and out-let suckers. Thus, under application of tubular turbulent apparatus of cylindrical design of jet type under technological design of fast chemical processes accompanied by local evolution of significant heat amount in real conditions there are several original ways of effective process heat regime control: change of reaction zone radius and rate of flow V, use of “zone“ model of realization of fast chemical process and application of shell-and-tube apparatus with N pipe bundle of small radius. Moreover tubular reactors should be applied in developed turbulent regime that allows both creation of homogeneous conditions in reaction zone by reagents concentrations (plug-flow regime in turbulent flows) and effective control of system heat regime. In section 1.3.3.4 we showed that it was possible to increase effectiveness of heat removal under fast chemical processes of polymer synthesis at the expense of intensification of convective heat-exchange (including intensification by preliminary reagents turbulization). Furthermore, profiling of heat-exchange surfaces of chemical reactors was widely spread with this aim. As a consequence, it seemed to be advisable to reveal the effectiveness of convective heat-exchange under proceeding of fast exothermal reactions in tubular turbulent apparatus in dependence on canal geometry and hydrodynamic structure of reaction mixture movement with the aim of development of regularization of heat regime at the expense of reaction zone geometry change.

4.2.2. Intensification of convention heat-transfer 4.2.2.1. Auuaratus neometrv effect Pressure difference at apparatus ends plays a significant role under selection of one or another construction of tubular turbulent apparatus under realization of fast exothermal chemical reactions. Consequently, investigation of influence of internal canal construction of tubular turbulent apparatus with external cooling on pressure difference under reaction mixture flow was carried out. Reaction zone of cylindrical design is characterized by the fact that values of pressure at the beginning and at the end are equal (Fig. 4.12). In this case low pressure difference at apparatus ends is observed and it does not exceed 0,03 atm at experimental conditions. Numerical dependences of pressure in tubular turbulent apparatus on volume flow of reaction mixture (R - correlation coefficient) were obtained: (apparatus beginning, R = 0,98) pb = 2.109~3947 pe = 2.10’0w3’86 (apparatus end, R = 0,99) (4.15) Ap = 4 . 1 0 w lgS (pressure difference Ap = pb - pe, R = 0,93) Significant increase of pressure at apparatus beginning in reaction zone of divergentconvergent design is observed (Fig. 4.13) and it determines high values of pressure difference (Ap reaches 1 atm). The change of pressure in tubular turbulent apparatus of divergent-convergent design on volume throughput of reaction mixture obeys the following regularities (R= 0,99): Ap E Pb = 1*105~2735 Pe = 2.1010w3t86 (4.16)

105

1 O03 A 1

80

100

120

140

160

180

200

220

w12683k.e~

Figure 4.12. Pressure difference (internal canal) in reactor of cylindrical design. 1 - beginning of apparatus (Pb); 2 - apparatus end (pe);3 - pressure difference (Ap).

1

30

50

70

90

110

13 150 w P , m 3kec

Figure 4.13. Pressure difference (internal canal) in reactor of divergent-convergent design. 1 apparatus beginning (pb); 2 - apparatus end (pe);3 - pressure difference (Ap). However, it is worth saying that according with (4.15) and (4.16) the pressure at end of apparatus of cylindrical and divergent-convergent designs obeys one law (Fig. 4.14). The pressures

106

at end of tubular turbulent apparatus of cylindrical and divergent-convergent designs are practically equal at comparable productivity. Thus, reactor of divergent-convergent design in comparison with cylindrical is characterized by high pressure difference (Ap increases in 25 times) that connected with high values of energy loss under reaction mixture flow through local hydrodynamic resistances. One should take it into account under realization of fast chemical processes and intensification of heat- and mass-transfer processes with the use of tubular turbulent apparatus in conditions of industrial polymer production. &,am

03

1

80

100

120

140

160

180

200

220

w a . m 3 bec

Figure 4.14. Pressure at the end of tubular turbulent apparatus of cylindrical (0) and divergentconvergent designs (*). Increase of resistance coefficient under reaction mixture flow in reactor of divergent-convergent design leads to the rise of heat-exchange effectiveness through the wall both at heating (Fig. 4.15) and at cooling (Fig. 4.16). In particular, at comparable reagents resistance times zr in apparatus of divergent-convergent and cylindrical designs in the first case temperature difference ATh along reactor length is in 1,5-1,7 times higher (Fig. 4.17). At the same time for reaching of one and the same value of AT1, liquid flows residence time in divergent-convergent canal in 1,8 times less in comparison with cylindrical apparatus. This determines the possibility of reduction of zr and consequently the possibility of proceeding of side reactions under realization of fast exothermal processes in tubular turbulent apparatus of divergent-convergent design. Since the heat flow Q at comparable temperature differences at apparatus ends ATh in divergent-convergent canal in 1,4 times higher (Fig. 4.18), so in this case increase of heating flow consumption (apparatus productivity) in 1,3-1,4 times is possible (Fig. 4.17). Thus, specific capacity of tubular turbulent apparatus of divergent-convergent design is in 1,3-1,4 times higher in comparison with cylindrical without heat load reduction that is a very favorable phenomenon allowing realization of energy-saving.

107 D Theat,OC

*l 6-

2

t

4.

2.

04 0

1

50

100

150

200 WE,

250

cm3/sec

Figure 4.15. Experimental dependence of internal flow heating AThat on volume flow w, (47°C) in reactor of cylindrical (1) and divergent-convergent ( 2 ) designs we = 130 cm3/sec (50°C).

Figure 4.16. Experimental dependence of reaction mixture cooling ATcoo]on volume flow w. (50°C) in reactor of cylindrical (1) and divergent-convergent (2) designs we = 40 crn3/sec (4-7°C).

108

12

3

.zoo 4

8 .

-

100

4.

rO

01

Figure 4.17. Dependence of heating of internal flow AT,,1 (1,2) and its volume throughput w. (4-7OC) (3, 4) on liquid flows (water) residence times Z~ in reactor of cylindrical (1, 3) and divergent-convergent design (2,4). we = 130 cm3/sec (60°C). Q ,Jlsec 2

1400

-

1000

i

600 0

100

200 WB.

cmJ/sec

Figure 4.18. Dependence of heat flow Q (heating) on flow rate in internal canal w. (4-7°C) of apparatus of cylindrical (1) and divergent-convergent (2) designs. we = 130 cm3/sec (60°C). It should be noted that effectiveness of heat-transfer through the wall under proceeding of fast chemical processes is increased at countercurrent way of reaction mixture and coolant flow p i g . 4.19).

109 DTCOOL

73

Om

75

77

79

81

83

a5

br, om

ei

Figure 4.19. Dependence of cooling of internal flow ATc,,,,l on temperature of internal divergentconvergent canal T, at forward flow (1-4) and countercurrent (1'-4'). w, = 36 cm3/sec, we (47OC) = 36 (1, 1'); 62 (2, 2'); 85 (3 , 3 ' ) ; 100 cm3/sec (4, 4'), dd = 26 mm; de = 10 mm; L, = 46 mm; L = 0,6 m. Using the equation of heat balance for calculation of heat-transfer processes heat-transfer coefficients K for studied apparatus (Fig. 4.20) were calculated with consideration of equation of liquid backflow (for divergent-convergent canal F = 0,044 m2, for cylindrical canal F = 0,05 m'). Calculations showed that in apparatus of divergent-convergent design heat-transfer coefficient was in 1,4-1,7 times higher in comparison with cylindrical canal and heat transfer effectiveness was reduced a little in transition region (Re = (4+10).103) (Fig. 4.20). 1,Jibl"cll

Figure 4.20. Dependence of heat-transfer coefficient K on the rate of internal flow WB and Reynolds' criterion Re in apparatus of cylindrical (1) and divergent-convergent (2) designs. It also seemed advisable to study dependence of heat-exchange efficiency on heat-carrying agent (coolant) volume throughput in annular canals of tubular turbulent apparatus of considered

110 3

constructions. It was found that at increase of annular flow rate Wa (w, = const = 130 cm /sec) in apparatus of divergent-convergent design increase of heat-transfer coefficient is observed (Fig. 4.21). At the same time in cylindrical apparatus leveling of Wa influence of effectiveness of convective heatexchange is observed. Since heat-transfer coefficient values are determined by lower heat-emission coefficient, so in apparatus of divergent-convergent design the possibility of effective intensification of heat-exchange is appeared at the expense of coolant throughput increase and this widens possibilities for heat-transfer processes intensification and control. 2

400

0

100

200 wa. crn3laec

Figure 4.21. Dependence of heat-transfer coefficient K on rate of coolant movement reactors of cylindrical (1) and divergent-convergent (2) designs.

Wa

in

Coefficient of heat-transfer through the wall determining intensity of heat-transfer processes depends on coefficient of heat-emission between the wall and the flow, that in its turn is the complex function of many factors, in particular liquid flow regime (laminar, turbulent), its heat properties, geometrical parameters of wall, etc.. Besides, heat-transfer coefficient K is significantly determined by value of lower heat-emission coefficient (a1or az).As a consequence, the increase of lower heatemission coefficient in accordance with (4.8) allows intensification of heat-exchange processes. In this connection, for revealing of dependence of heat-transfer (heat-emission) coefficient and consequently of the effectiveness of heat-transfer processes on hydrodynamic structure of reaction mixture and coolant movement in tubular turbulent apparatus distributions on flow particles residence times in heat-exchange zone were investigated (see 2.2.8).

4.2.2.2. The correlation between heat-transfer coefficientand hydro-dynamic regime o f reaction medium motion Experimental results showed (Fig. 4.22) that reactor of divergent-convergent design is characterized by low values of Bodenstain criterion Bo = 25i45 (high rate of longitudinal mixing E = (14+40)10-3 m2/sec (Fig. 4.23)), intensive heat-exchange is corresponded to this (IC = 600+650 J/(m2.sec.K)) (Fig. 4.20). Besides, due to high rate of longitudinal mixing (turbulence) in divergentconvergent canal formation of quasi-isothermal (isothermal in any cross section of apparatus) regime determining effectiveness of external heat removal is possible, that is necessary to take into account under apparatus design of fast polymerization processes. Cylindrical construction of tubular apparatus is characterized by high values of criterion B o = 60+100 (Fig. 4.22) (low rate of longitudinal mixing E = (2+10)10-3 m2/sec (Fig. 4.23)) and consequently by comparatively low effectiveness of heat-exchange (IC = 350+460 J/m2,sec.K) (Fig. 4.20). Thus, heat-transfer (heatemission) coefficient in spite of factors mentioned above is also a function of structure of reaction

111

mixture movement characterized by criterion Bo. This determines possibility of use of reception method of curves of response to introduction of inert indicator for modelling of heat-exchange processes in industrial reactors. In particular, under increase of volume flow of reaction mixture wv in divergent-convergent canal the values of criterion Bo are decreased (Fig. 4.22) that determines the rise of heat-transfer coefficient (Fig. 4.20). At the same time for cylindrical canal dependence of criterion Bo on liquid flows rates wv has distinct extreme character (Fig. 4.22). In transition region, in particular at Re = (2,5+8)103 the rise of criterion Bo with wv increase is observed that approaches hydrodynamic structure of flows to plug-flow regime. This determines comparatively low effectiveness of heat-exchange in cylindrical canals in this region of Re values (Fig. 4.11, 4.20). According with results presented in Figures 4.20 and 4.22 it is obvious that for heat-transfer coefficient increase and consequently rise of effectiveness of heat-exchange processes in cylindrical canal use of high-speed liquid reagents flows is required, that however requires increase of heatexchange surface (reaction zone dimensions). BO

1

lZO

Figure 4.22. Dependence of Bodenstain criterion on internal flow rate wv and Reynolds' criterion Re in reactors of cylindrical (1) and divergent-convergent (2) designs. Rather different picture is observed under formation of hydrodynamic flows (coolant) structure in annular canals of tubular turbulent apparatus (Fig. 4.24). Flows structures in annular canals of cylindrical and divergent-convergent apparatus (Bo = 80) practically coincide up to volumetric flow of liquid flows wa = 110 cm3/sec. At wa > 110 cm3/sec in divergent-convergent apparatus reduction of criterion Bo is observed that is determined by the rise of longitudinal mixing rate (Fig. 4.25) and in cylindrical canal the rise of criterion Bo is observed (Fig. 4.24).

112

E 103. m2/sec

".

,

70

30 2.4

5.6

cm3/sec

150

110

12

8,8

Re 10-3

Figure 4.23. Dependence of longitudinal mixing coefficient E on reaction mixture movement rate wv and Reynolds' criterion Re in reactor of cylindrical (1) and divergent-convergent (2) designs. Bo

100

-

80-

0

0 2

60

.-

-

0

. 30

60

90

120

150 wa, sm3/sec

Figure 4.24. Dependence of Bodenstain's criterion on coolant movement rate wa in cylindrical (1) and divergent-convergent (2) apparatus.

113

E 103,m 2 h c

0 2

1

4

0.3 30

60

90

120

150

wa, cm3Isec

Figure 4.25. Dependence of longitudinal mixing coefficient E on coolant movement rate cylindrical (1) and divergent-convergent (2) apparatus.

Wa

in

One may suggest that for increase of heat-emission coefficient of flow a2 (coolant) in annular canal of divergent-convergent apparatus and consequently heat-exchange effectiveness the rise of movement rate Wa is advisable especially when a2 < a1 because IC = a2.Tubular turbulent apparatus with cylindrical canal is quite on the contrary inexpediently to exploit at high rates of heat-agent flow in annular canal. Really, at annular flow rate increase Wa (w, = const= 130 cm3/sec) in divergentconvergent apparatus the rise of heat-transfer coefficient is observed (Fig. 4.21). At the same time for cylindrical apparatus leveling of Wa influence on effectiveness of convective heat-exchange is characterized (Fig. 4.21). Thus, from the point of view of heat-exchange effectiveness necessity of deep profiling of reaction zone in the form of divergent-convergent construction under realization of fast exothermal reactions and heat-exchange apparatus under polymers synthesis was confirmed. Divergent-convergent construction without change of heat-exchange effectiveness allows increasing of specific apparatus productivity in 1,3-1,4 times at heat-agent (reaction mixture) residence time decrease in 1,8 times. Dependence of heat-emission coefficient of liquid flows on hydrodynamic structure of rection mixture movement determined by Bodenstain's criterion Bo was revealed. Low effectiveness of heat-exchange processes in cylindrical canal in comparison with divergent-convergent in tubular apparatus is determined by high approximation degree of structure of flows moving both in internal and annular canals to plug-flow regime that is undesirable. In apparatus of divergent-convergentdesign effectiveness of heat exchange at a2 < a1 may be considerably intensified by increase of coolant movement rate Wa that widens possibilities for heat-transfer processes intensification and control. The following conclusions may be drawn out of results of convective heat-exchange investigation: heat-transfer coefficient in tubular turbulent apparatus of divergent-convergent design is in 1,41,7 times and specific productivity is in 1,3-1,4 times higher in comparison with tubular apparatus of cylindrical design under decrease of reaction mixture residence time in heatexchange zone in 1,8 times; correlation between heat-transfer coefficient and liquid flows hydrodynamic structure is revealed. High degree of approximation to plug-flow regime of mixing and high values of heat-

114

transfer coefficient correspond to low Bodenstain's criterion values that are predominantly realized in divergent-convergent canals. Fundamental and applied aspects of fast exothermal reaction proceeding worked out in presented chapter including polymerization allow solution of big problem of temperature regime control in reaction zone at the expense of external heat removal. Formation in reaction zone of quasi-plug-flow mode in turbulent flows when bounds of "torch" reach heat-exchange surfaces determines possibility of thermostating of fast chemical reactions at the expense of external heat removal. Numerical dependences obtained in the work allow assured formation of this regime at conditions of industrial production in dependence on kinetic parameters of chemical reaction and physical characteristics of reagents flows. Suggestion on nature of macroscopic fronts formation in turbulent regime under mixing of reacting and neutral flows was made. Formulas fitting for engineering calculation of temperature profile in reaction zone at external thermostating were obtained. On their base technologically suitable ways of heat regime control were revealed, in particular by change of reaction zone geometry, hydrodynamic regime of reaction mixture movement, apparatus design. Qualitative correlation of heat-transfer coefficient with hydrodynamic structure of reaction mixture movement at convective heat exchange was revealed.

115

CHAPTER 5. NOVEL SOLUTIONS IN THE FIELD OF FAST CHEMICAL PROCESSES REALIZATION IN TURBULENT REGIME UNDER POLYMERS SYNTHESIS Conception of tubular turbulent apparatus application is developed completely enough on the base of macro-kinetic approach to investigation of isobutylene electrophilic polymerization. At the same time results received during studying of regularities of turbulent mixing in homogeneous and heterogeneous systems and also temperature regime control in reaction zone under fast exothermal processes that were presented in Chapters 3 and 4 allowed significant broadening of tubular turbulent apparatus application field under polymers synthesis. In particular, possibility of scientifically-grounded application of novel technologies both for new chemical processes class fast polymerization reactions and for intensification of heat- and mass-transfer was appeared. This determines advisability of working out of novel solutions in the field of fast chemical processes realization in turbulent regime and consequently creation of continuous energy- and resource-saving technologies on the base of tubular turbulent apparatus under synthesis of other polymer products in particular those considered in 1.4.

5.1. Butyl rubber halogenation Numerical value of chemical reaction characteristic time is of fundamental importance in respect to possibility of realization of novel continuous process of chlorobutyl rubber synthesis. According with [27] in temperature interval 290-325 K time of chlorination reaction is less than 60 sec, and in concrete case of chlorination of 1 5 1 6 % butyl rubber solution in methylchloride (328 K, dosage of molecular chlorine - 3 - 3 3 mass %) is equal to 7,5f2,5 sec [176]. Close values of chemical reaction characteristic times were obtained also under experimental investigation of chlorobutyl rubber reception by reaction of butyl rubber interaction with chlorine in nefrase solution in turbulent flows. Calculations of geometry of butyl rubber chlorination reaction zone were carried out on the base of regularities of turbulent mixing in self-similar regime (see Chapter 3). At rate of introduction into tubular apparatus of butyl rubber solution in nefras e ual to 0,21 m3/h and nitrogenous chlorine mixture (5 : 1 volumetric) introduction rate equal to 2 , l m /h the linear reagents movement rate without consideration of occupied by special fillings volume come to V = 0,33 d s e c and consequently residence time is Z~ = L / V = 6,l sec. Taking into account the fact that there was no slip of molecular chlorine after reaction mixture output from apparatus in the course of device functioning we supposed that chlorine conversion was not less that 99% and chemical reaction time Tch was comparable with reaction mixture residence time in apparatus z ~ . Since chemical reaction time is small ( ~ ~ 1=, 6 sec) liquid-phase butyl rubber chlorination process in solution with molecular chlorine should be ascribed to novel class of chemical processes - fast chemical reactions that should be carried out also according with principally novel technology with the use of highly productive compact tubular turbulent apparatus of jet type. Experimental-industrial tests showed that butyl rubber chlorination in solution by molecular chlorine in turbulent flows is provided under application of tubular turbulent apparatus of jet type of divergent-convergent design as reactor-chlorinator, homogeneous gaseous chlorine distribution along reaction volume that guarantee formation of polymer stable in its structure (Table 5.1). For example, dispersion by chlorine content in CBR received in turbulent flows was not higher than 0,05 mass %. At the same time for samples received by standard technology in stirred tank reactors this index was 0,14 mass %, i.e. homogeneity of resulted CBR samples was increased practically in 3 times.

93

11G

Table 5.1.

reaction zone (mixer) It is obvious from obtained results (Table 5.1) that brominated butyl rubber can be synthesized by proposed method. Moreover in turbulent regime halogenation of other elastomers in particular ethylene and propylene copolymers was carried out. Experimental-industrial tests showed that in spite of high exothermicity of butyl rubber chlorination (q = 184 kJ/mole) heat regime in zone of reaction proceeding didn't course any problems that is confirmed by calculations. Assuming that values of average heat capacity and reaction mixture density in the first approximation are equal to the same parameters of solvent with temperature adiabatic rise in reaction zone according with (1.4) at productivity of experimental device 23 kg/h is AT = qAn/ C,$ = 184.1,6 / 0,17460 = 2°C. Thus, even at adiabatic regime of tubular turbulent apparatus-chlorinator work (without heat removal) temperature rise in reaction zone under butyl rubber chlorination (12- 15% solution) by molecular chlorine in tubular apparatus functioning in optimal regime of ideal plug-flow in turbulent flows doesn't exceed 2fl"C. One may consider that process proceeds in quasi-isothermal conditions and doesn't require external or internal heat removal and special mixers. It should be noted that in technological scheme of CBR production tubular turbulent apparatus of jet type of analogous construction should be used also at other stages of technological process, in particular under neutralization of resulted CBR solution (constant of rate of interaction between mineral acids and alkalis is about k = lo9" Vmole.sec), washing out of CBR solution from salts, etc. (extraction), purification of inverse solvent (extraction), under introduction into CBR solvent of stabilizer-antioxidant and anti-aglomerator (mixing) instead of intensive including volumefree agitators with mechanical mixers. Thus, in chlorobutyl rubber production stages of butyl rubber chlorination in solution by molecular chlorine, neutralization and polymerizate washing out and also stabilizer introduction into CBR solution are related to fast chemical and mass-exchange physical processes that should be carried out in turbulent flows. Possibility of realization in tubular turbulent apparatus of self-similar regime in relation to viscosity and also revealed regularities of multi-phase flows allow calculation of optimal apparatus construction under creation of continuous energy- and recourse-saving technology of chlorobutyl rubber production. Moreover obtained results determine possibility of application of novel technology for both fast polymerization processes and fast polymer-analogous reactions of polymers modification.

5.2. Cationic piperylene oligomerization Electrophilic catalysts that are used for cationic piperylene oligomerization posses various activity (see 1.4.2). This determines principle possibility of improvement of industrial production of synthetic drying oil at stage of piperylene cationic oligomerization by selection of catalytic systems with corresponding activity and selectivity.

1 I7

5.2.1. Kinetics of piperylene cationic oligomerimtion Oligomerization reaction for all considered electrophilic catalysts is characterized by the first orders by monomer and catalyst. In particular, this fact is shown on the example of Tic14 and A1C13*O(C6H5)2 use (Fig. 5.1). As it is obvious activity of studied catalysts is changed in wide diapason (Table 5.2) and piperylene oligomerization process in the presence of Al(C2H5)zCl and Al(i-C4H9)2Cl practically does not proceed. High yields (7040% and higher) can be reached for 1520 sec if AlC13.O(C6Hs)2 and also AlCzHsCl2 are used as a catalysts. And in the presence of Tic14 the product yield is not higher than 0,5-1% for the same time (Fig. 5.2). Addition of electron-donor compounds, in particular (C&)20 to the electrophilic aluminum-containing catalyst leads to reduction of reaction rate constant at piperylene oligomerization that is obviously coursed by catalytic complex acidic force decrease. They may suppose that AN213 will reveal higher activity in comparison with AlC13.o(C6H5)2 on the base of principle of correlation between acidity and cationic activity of Fridel-Krafts catalysts in polymerization processes. However in this case due to catalyst's high activity in electrophilic processes cross-linked polymer will be formed [246] that is inadmissible. 0

0,02

0,04

0,06

Cc, molell 0,08

2

.-C

E

T

6 %

E

P 4

2

a, mole11 0

3

6

9

Figure 5.1. Dependence of rate (V,) of piperylene oligomerization initiated by AIC1343(C6H& (1,2) and Tic14 (3,4) on monomer (1, 3) and catalyst (2,4) concentrations. 1 - CAI= 0,014 mole/l; 2 , 4 - Cm = 1 mole/l; 3 - CTi = 0,025 mole/l; (To= 298 K).

118

80

-

60

-

40

-

5

20.

0

5

10

15

20

25

30 tp, sec

Figure 5.2. Conversion curves of piperylene oligomerization initiated by various catalysts Tic14 (I), TiCkAl(i-C4H9)3 ( 2 ) , A ~ C ~ H ~ C ~ ~ . O ((3), C~H AlCzHsC12 S)~ (4), AlC13.O(CsH5)2 (5). CTi = 0,025 mOle/l (1, 2); c A l = 0,02 (3, 4) and 0,028 (5) mole/l; CM= 1,o (1, 3-5) and 2,0 (2) mole/l. T = 298 K. As in the case of cationic polymerization of others diene hydrocarbons oligopiperylene macromolecules synthesized in the presence of studied catalysts contain predominantly rrans-units (Table 5.2). Increased content of cis-1,4-units is observed for samples received in the presence of TiCkAl(i-C4H9)3 and AlC2HsC12. Analysis of double bonds content in oligopiperylene macromolecules showed that for all catalysts applied in synthesis process insignificant change of unsaturation in forming polymer products was observed (Table 5.2). Addition into the catalyst of ether as modifier, in particular of (C6H5)20 in ratio 1 : 1 mole / mole reduces both electrophilic catalyst activity (kpis decreased) and possibility of gel-formation processes proceeding (cross-linking). Molecular-mass distribution of oligopiperylene obtained on studied catalysts except TiC14Al(i-C4H9)3 system is close to possible (Table 5.2) and is unimodal (Fig. 5.3). In the presence of TiC14-Al(i-C4H9)3 the curve of distribution of oligomer MM displaces to the high-molecular region (curve shape takes well-defined bimodal character). As a consequence MMD is broadened that can be explained by the presence of several active sites types in the catalyst, in particular both cationic and ion-coordinated responsible for obtaining of mixture of low- and high-molecular polymers correspondingly. Oligopiperylene formation with such molecular characteristics is obviously technologically unprofitable due to polimerizate viscosity increase at the expense of product with high molecular mass and worsening of synthetic drying oil operational characteristics.

119

Table 5.2. Cationic oligomerization of piperylene initiated by some electrophilic catalysts (conditions as in

Thus, using accessible method of process kinetics control the row characterizing kinetic activity of prevalent electrophilic catalytic systems for piperylene cationic oligomerization was received: Tic4 < TiCL-Al(i-C4H9)3< AIC2HsC12O(C6Hs)2 < AICzHsC12 < AlCb.O(C6Hs)z. Studied catalysts are characterized by broad, differing in to orders activity range (k,, = 21420 Vmole.min), without significant differences in molecular characteristics of resulted oligomer products. Application of two-component Ziegler-Natta catalytic system leads to reception of polymers mixture with increased average MM and broad MMD of oligopiperylene at insignificant increase of cis-l,4-units content in molecules. Cationic piperylene oligomerization under liquid synthetic oligopiperylene rubber production in the presence of various catalytic systems according with obtained experimental data (kinetic parameters) required various approach to technological process design. In particular, high enough value of constant of piperylene oligomerization rate in the presence of AlCI3.O(CsH5)2 (kp= 420 I/mole.min) and AlC2H5C12 (k,, = 200 I/mole.min) is paid attention. In this case the use of AlC13,0(C6H5)2 as piperylene oligomerization catalyst allows significant simplification of technological scheme of synthetic drying oil production and creation of energy- and resource-saving process of high ecological purity with the use of compact tubular turbulent reactors functioning in plug-flow regime in turbulent flows.

120

160

120

80

40

0 0

1

2

3

Figure 5.3. Curves of molecular-mass distribution of oligopiperylene. AICZH~C~Z~(C~H AlCzHsClz ~)T Conditions as in Figure 5 . 2 .

-TiC14....msaTiCl4-Al(i-C4Hg)3~-

5

4

--

-

A~CI~.O(C~H~)Z

5.2.2. The influence of kinetic parameters on polymerization reactor selection The range of chain propagation reaction rate constants k, differs practically in two orders without essential differences in molecular characteristics (Mw, M,, M,/M,) of synthesized oligopiperylene (Table 5 . 2 ) . This determines necessity of use in technological scheme of liquid oligopiperylene rubber synthesis of chemical reactors constructions of various types in dependence on catalytic systems activity. The key parameters for selection of reactor type are residence time and reaction zone dimensions that are necessary for achievement of required monomer conversion in isothermal conditions and consequently the optimal time of process carrying out. The values of effective rates constants of piperylene polymerization initiated by various catalytic systems and required residence times are presented in Table 5.3.. Obviously, for piperylene oligomerization on Tic14 and TiC14-Al(i-C4H9)3 catalysts (k,, = 0,03-0,15 I/mole.sec at 298 K) standard stirred tank reactor or cascade of such reactors is advisable to use because of big times of reaction mixture residence zr in apparatus. In the case of oligopiperylene synthesis in the presence of AlC13.O(C6H5)2, AlCzHsClz and possibly A ~ C Z H ~ C I Z . O ( Chigh ~ H ~values ) ~ of rate constants (kp= 1-7 Vmolessec at 298 K) and correspondingly small geometrical reaction zone sizes determine necessity of use of compact tubular turbulent apparatus that can operate in quasi-plug flow regime (see 4.1). Moreover, low values of zr (2-20 sec) in this case reduce possibility of proceeding of side extremely undesirable processes (cross-linking, macromolecules cyclization, etc.). Since quasi-plug flow regime in turbulent flows is formed at definitely limited reaction zone sizes, so it is necessary to consider the ratio of characteristic times of mixing zmix and chemical

121

reaction rch under piperylene oligomerization in the presence of studied catalysts with the aim of selection of optimal reactor construction. Table 5.3. Piperylene oligomerization in stirred tank reactor and tubular plug flow a paratus initiated by various catalysts (x = 0,95, w = 4,4 m /h, Cc = 0,15 mole/l, C,O = 5,7 mole/l)

P

Tubular turbulent

Proceeding of oligomerization process at the absence of diffusion kimitations in reaction zone of divergent-convergent design at self-similar reaction mixture flow regime is determined by condition (see 3.1.4.1):

here Rd - radius of wide reaction zone part where homogeneous field of reagents concentrations or required reactor (diffusor) radius is necessary to create (Fig. 3.lb); w - reaction mixture volume throughput; Cc- catalyst concentration; x - monomer conversion degree; fE, f, and f - numerical coefficients calculated from (3.15), (3.22) and (3.23); Ls- the length of divergent-convergent section with openning angle y; & - radius of narrow (convergent) part of reaction zone (Fig. 3.1,b). Using (5.1) we may obtain the equation for calculation of critical radius of reaction zone RCr d that is the criterion of quasi-plug flow regime formation at achieving of following correlation Tnlix 5 zch (Ls/ 2& = 1,7, & / & = 1,6, y = 45'):

The values of R Y (Table 5.3, Fig. 5.4) for piperylene oligomerization reaction fast enough initiated by catalysts A1C13.O(C&)2 and AICzHsC12 confirm possibility and necessity of use in this case of tubular turbulent apparatus where intensive turbulent mixing of reaction mixture is reached at the expense of hydrodynamic flow energy. Low reaction rate under piperylene oligomerization on catalysts Tic14 and TiCkAl(i-C4H9)3 (Table 5.3) determines advisability of use of standard stirred tank reactors (Fig. 5.4). Use of active catalyst A1C13.o(C&)2 and realization of piperylene oligomerization process in tubular turbulent apparatus in comparison with stirred tank reactor and TiCL usage allows decreasing of reagents residence time in reaction zone zr and its volume Vr in 1326 and 1343 times correspondingly.

122 Rdcr, rn 0,16

0,12

0,OB

0,04

0

2

1

0

3

4

5

kp, I/&Ie sec

*

Figure 5.4. Dependence of reaction zone radius R Y on kinetic parameters of piperylene oligomerization in the presence of various catalytic systems (conditiobs as in Fig. 5.3). When they select reactor for any chemical process realization including polymerization they should take into account temperature distribution in reaction zone and possibility of heat regime control with the aim of receiving of polymer with close to calculating Mh4 and h4MD. Temperature field in apparatus with reaction zone geometry calculating in accordance with kinetic parameters of piperylene oligomerization process by (5.2) (Table 5.3) at adiabatic conditions of process proceeding is changed in accordance with (see chapter 4): T=To+Tad =To+

(5.3)

cPp With reference to piperylene polymerization transforming (4.9) we may obtain correlation for calculation of temperature along reaction zone length L at external heat removal: m

m

.

m

,m

m

.

m

\f .

,

L

For studied catalytic systems temperature rise in reaction zone exceeds maximum permissible for piperylene oligomerization (393 K) even at external heat removal (Fig. 5.5). Reduction of temperature, i.e. heat removal in reaction zone at the expense of cooling in the case of A ~ C I ~ . O ( C ~ Hand S ) ~AlC2HsC12 catalysts use, when reaction rate is high enough is practically observed after chemical process completion (Fig. 5.5, curves 6, 7). This determines low effectiveness of external heat removal at the tank reactor use when reaction zone does not reach heat-exchanging surfaces. At the same time significant temperature reduction in reactor can be reached under tubular turbulent apparatus application functioning in quasi-plug-flow mode under piperylene oligomerization in the presence of AlC13.0(CsH5)2 at the expense of small reaction zone

123

radius Rd = 0,022 m (Fig. 5.5, curve 6). Only in this case the optimal regime of piperylene oligomerization fast process proceeding will be realized in jet regime on corresponding catalysts without additional mechanical mixing. Thus, conception of fast chemical reactions realization in turbulent flows limited by impenetrable wall is applicable not only for separate class of process always proceeding with high rates. Differences in kinetic parameters under proceeding of one and the same chemical reaction in the presence of catalysts differing in activity determine the necessity of various approach to selection of type and construction of main reactor. The preference obviously should be given to the most active catalytic system that allow creation of high-productive, energy- and resource-saving technologies 011the base of compact tubular turbulent apparatus. Possibility of guarantee formation in reaction zone of quasi-plug-flow mode in turbulent flows (see 4.1) and also application of divergent-convergent construction (see 4.2.2) allow solution of heat regime control problem under piperylene polymerization process proceeding in the presence of A1C13.0(C6H5)z. Low values of dynamic viscosities of liquid oligopiperylene rubber at reactor outlet (at 283 K p = 1 mPa.sec) allow recommending of compact tubular turbulent apparatus use in this process also at stage of catalyst deactivation by propylene oxide or water.

570

1

470

370

270

0

4

8

L,m

Figure 5.5. Temperature change in reaction zone along tubular reactor under piperylene oligomerization at adiabatic conditions (1-5) and at external heat removal (6-10). Catalysts Tic14 (1, lo), TiC14-Al(i-C4H9)3 (2, 91, AIC2HsC12.0(CsHs)z (3, 8), AlCzHsC12 (4, 7), AlC13.O(C6H5)2 ( 5 , 6 ) . To = 273 K, Tch = 273 K, C,O = 5,7 mole/l, C, = 0,15 mole/l, R d = R Y -Table 5.3.

5.3. Separation of fast and slow stages under polymers synthesis in the presence of ZieglerNatta catalysts Polymerization process consists of several consecutive stages. Initiation, chain propagation and termination are the main of them. Optimal conditions are necessary for proceeding of corresponding stages to produce polymer product of high quality. These conditions conclude

124

working temperatures, pressure, reagents concentrations and sufficient degree of homogeneity of all characteristics fields. At the same time, we can observe situations when the main processes such as chain propagation proceed slowly and previous stages are fast. Both chemical (initiation of olefins and dienes (co)polymerization in the presence of Ziegler-Natta catalysts) and mass-transfer physical (saturation of solvent by gaseous monomers under ethylene and propylene copolymerization) reactions may present fast stages in gross-process. Fast processes as we have shown in part 1.1.4 proceed in diffusive regime. Consequently we observe irregular distribution of reaction mixture components both on micro- and macrolevels when we introduce reagents directly into the stirred tank reactor. This fact determines irregularities in macromolecules structures, alteration of molecular characteristics of resulting polymers and consequently deterioration of their quality (see 1.4.1 and 1.4.3). It turned out to be expedient to separate fast stages (chemical and mass-transfer physical) from properly (co)polymerization of olefins and dienes on Ziegler-Natta catalytic systems. In this case for every stage it is possible to provide ideal (or nearly ideal) conditions for proceeding of corresponding processes. For the first fast stage it can be attained in turbulent pre-reactor of divergent-convergent design, and for the second - in tank reactor-polymerizer. While using pre-reactor of divergent-convergent design it is necessary to obtain ratios zfix I zC1,(when chemical process proceeds at fast stage) and zfix I zch (when mass-transfer process proceeds at fast stage) with the help of regularities of turbulent level increasing (see chapter 3). Separation of stages is considered by the example of isoprene polymerization and ethylene and propylene copolymerization in the presence of Ti-AI and V-AI catalytic systems, accordingly.

5.3.1.Stereospecific polymerization of isoprene The untraditional way to influence on molecular characteristics of resulting polymer product under dienes polymerization initiated by Ti-A1 catalysts of Ziegler-Natta type was exposed. The observed effect is reached by separation of fast stage of formation of sites on which macromolecules propagate from the slow stage of polymerization. Fast stage was carried out under conditions of high turbulence in reaction zone with the use of compact tubular turbulence apparatus. And polymerization stage was carried out in traditional stirred tank reactor. Because of Ti-AI catalytic systems are micro-heterogenious it is interesting to investigate the manner of moving of solid catalyst phase particles in the reaction zone of divergent-convergent type, the last ones as it was mentioned above in part 3.2.4 were characterized by the appearance of separation effect. According with literature data [174] at the initial moment of mixing of Tic14 and Al(i-C4H9)3 solutions particles as thin plan polygons with medium diameter d2 about (0,3+1).10-’m are formed. If these particles are kept in the absence of monomer they stick together and form secondary larger aggregates with diameters about 30 mcm. If we take mentioned sizes as a diameter of particles of catalyst disperse insertions we may get evaluation values of phases separation effect for Ti-A1 catalytic system when reaction medium flows in a tubular turbulent apparatus. Calculations on the base of (3.48) have shown (Fig. 5.6) that when catalyst particles diameter d2 and linear rate of reaction medium movement V, increase under isoprene polymerization on Ti-A1 catalytic system relative difference of volume fractions of disperse phase (of catalyst) at device’s axis and its periphery also rises. However, separation effect is slightly displayed ACX~ / CX~,,, - lo-’ because particles of heterogeneous catalyst are of very fine sizes. Thus, under the use of tubular turbulent pre-reactor of divergent-convergent design for carrying out of fast stages of formation of sites of macromolecules propagation and initiation under isoprene polymerization on Ti-A1 Ziegler-Natta catalysts homogeneous spreading of particles of heterogeneous catalyst upon the whole pre-reactor volume is provided. In this case similar conditions for movement of all potential active sites in reaction zone of divergent-convergent design are provided.

125

Hydrodynamic effect on solution of separately prepared Ti-AI catalytic system in turbulent regime (variant 1, see 2.2.2) practically doesn't change isoprene polymerization rate in comparison with traditional way of process carrying out (free experiment) (Fig. 5.7, curve 1,2). In this case catalyst can be activated at the expense of catalytic system formation by mixing of initial components directly in tubular turbulent pre-reactor with further aging at 273 K (variant 2, see 2.2.2) (Fig. 5.7, curve 3). Some other picture in conversion curves is observed under modification of Ti-AI catalytic system by dienes additives (piperylene). Under the use of modified Ti-A1 catalyst isoprene polymerization rate is significantly increased even at free experiments (Fig. 5.7, curves 1, 4)that is widely used in industrial production for activation of Ziegler-Natta catalytic system under dienes polymerization. Hydrodynamic effect on triple catalytic system under its mixing with solvent in turbulent regime (variant 1) allows at activity maximum by free experiments (Fig. 5.7, curves 1, 4) at polymerization time 1 hour increasing of polymer product yield practically for 10% (Fig. 5.7, curve 5). Hydrodynamic effect on reaction mixture at the moment of its formation is more radical in relation to isoprene polymerization rate, i.e. under mixing of monomer and catalytic complex solutions in turbulent regime (variant 3, see 2.2.2). In this case high polyisoprene yields (-80 %) can be already reached at polymerization time 20 min (Fig. 5.7, curve 6). Obviously, isoprene polymerization rate increase in considered examples is connected with breaking of particles of micro-heterogeneous catalysts on the base of TiCl3. Dispersing analysis of particles of received catalytic systems was carried out with the aim of this suggestion adequacy confirmation.

Figure 5.6. Dependence of relative difference of inclusion volume fraction of TiC14-Al(i-C4H9)3 catalyst particles on the axis of tubular turbulent pre-reactor of divergent-convergent design and its peripheral part Aa2lazn, on particles diameter d2 and linear rate of reaction mixture movement V, under isoprene polymerization. P I = 867 kg/m3, p2 = 2670 kg/m3, P I = 0,552 mPa.sec.

126

Y$H,% 6

0

20

40

60

tp , m i l

Figure 5.7. Isoprene polymerization in the presence of TiCl4-Al(i-C4H9)3 (1-3) and TiC&-Al(iC4H9)3-piperylene (4-6) catalytic system. Variant (see 2.2.2) 1 (2, 5 ) ; 2 (3); 3 (6); free experiment (1, 4). CTi = 6 mmole/l, c~= 1,5 mole/l, Ti / A1 / piperylene = 1 / 1,25 / 2, zvCc = 30 min at 273 K, solvent - toluene, 298 K. Under Ti-A1 catalytic system formation in the absence of diene modificating additives catalysts particles large enough are formed (the average radius is about 304 mcm). Hydrodynamic effect on two-component Ti-A1 catalytic system in turbulent regime (variant 1) in comparison with traditional way of process carrying out practically does not influence on catalyst's particles sizes and consequently its specific surface (Fig. 5.8, curves 1, 2) that determines equality of polymerization rates (Fig. 5.7, curves 1, 2). Formation of Ti-A1 catalytic system in the presence of piperylene additives (ternary catalytic system) leads to decrease of catalyst's particles down to re = 1,5 mcm (Fig. 5.8, curve 3). Since active sites of stereospecific polymerization are placed on defects of catalyst's particles crystal structure, in given case - P-TiCls, so increase of its specific surface may lead to the polymerization rate rise that was experimentally observed (Fig. 5.7, curve 4). Hydrodynamic effect on modified TiAI catalytic system in turbulent regime (variant l ) leads to the additional decrease of catalyst's particles sizes down to re = 1 mcm (Fig. 5.8, curve 4) that determines rise of polymerization rate in comparison with traditional way of process carrying out (Fig. 5.7., curves 4, 5). Probably under catalyst formation in the presence of diene on the active particles propagation of macro-chains consisting of piperylene begins, that leads to weakening of morphological structure of solid phase and lighten its breaking under the action of energy of turbulent flow. Hydrodynamic effect on such depositions at the expense of use of tubular turbulent apparatus (variant 1) is adequate to fragmentation of heterogeneous Ziegler-Natta catalysts in polymerization process. Obviously polymerization rate rise under catalytic system formation in turbulent regime (variant 2) (Fig. 5.7, curve 3) when mixing of Tic14 and Al(i-C4H9)3 solutions occurs directly in tubular turbulent prereactor, is also explained by increase of specific catalyst surface. In this case at high rate of micromixing creating in turbulent pre-reactor the nucleation rate-to-crystal gross rate ratio is increased that leads to increase of nucleus number of new phase and to decrease of re.

127

1

150

100

50

0

n

2

4

8

6

re.man

Figure 5.8. Differential curves of particles distributions of TiC14-Al(i-C4H9)3 (1, 2) and TiCLAl(i-C4Hg)3-piperylene (3,4) catalytic systems by sizes. Variant (see 2.2.2) 1 (2,4); free experiment (1, 3). C T=~ 12 mmole/l, Ti / AI / piperylene = 1 / 1,25 / 2, zyCc = 30 min at 273 K, 298 K. Experimental results showed that increase of catalyst specific surface under stereospecific dienes polymerization is evident but not sufficient way of process proceeding rate increase. Significant activity of catalytic system is revealed at preliminary mixing of catalyst and monomer solutions mixing in turbulent regime (variant 3) (Fig. 5.7, curve 6). In this case fragmentation of catalyst's particles occurs (Fig. 5.8, curve 4) and its uniform spreading upon reaction mixture. Moreover, effect intensification is possible at the expense of diffusion limitations decrease under the first monomer molecule connection to the active site (initiation) that proceeds with high rate. Thus, by changing of way of isoprene polymerization process carrying out in the presence of Ziegler-Natta catalytic systems namely by applying of tubular turbulent pre-reactor containing no additional mixing and active devices with reaction mixture residence time in mixing zone about 2-3 sec one may effectively influence on process proceeding rate. Under modifying of Ti-AI catalytic systems by diene additives increase of specific area of catalyst and consequently the rise of polymerization rate occur. This effect is intensified at the expense of additional fragmentation of catalyst's particles under hydrodynamic influence on catalytic system in turbulent regime. Polymerization rate rise is maximum under preliminary monomer and catalyst mixing in turbulent regime, i.e. under increase of specific area of catalyst and uniform distribution of its particles in reaction mixture. Considered methods of isoprene polymerization carrying out in the presence of micro-heterogeneous Ti-AI catalytic system allow decreasing of catalyst consumption and consequently prime cost of resulted product. Separation of fast stage of active sites formation and initiation with increase of turbulent mixing level at this stage at the expense of tubular turbulent pre-reactor use leads to the change of molecular characteristics of synthesized polyisoprene. In particular hydrodynamic effect on separately prepared TiC14-AI(C4H9)3 catalytic system (variant 1) leads to the rise of both weightaverage M, (Fig. 5.9) and numerical-average Mn (Fig. 5.10) molecular masses. In spite of the fact

128

that stage separation in this case does not lead to the change of polymerization rate (Fig. 5.7, curves 1.2). MwlOd

e 0,8

0 $6

-

e e/

OA

0.2

i

0

20

40

60

100

80

@,mil

Figure 5.9. Dependence of M, on isoprene polymerization time T, in the presence of TiCLAl(iC4H9)3. 1 - free experiment, 2 - hydrodynamic effect on catalytic system in turbulent regime (variant 1, conditions as in Figure 5.7). In the case of isoprene polymerization initiated by modified ternary TiCL+-AI(i-C4H9)3-piperylene catalytic system at separation of fast and slow stages the rises of both polymerization rate (Fig. 5.7, curve 5,6) and molecular masses (Fig. 5.11,5.12) are observed. Analogous increase of average molecular masses was observed also under formation of catalytic system (variant 2) and reaction mixture (variant 3) in turbulent regime). Change of polymerization carrying out method influences on the width of molecular-mass distribution (Fig. 5.13). In the case of TiCl~-Al(i-C4H9)3increase of turbulence level at active sites formation stage leads to MMD broadening (Fig. 5.13, curve 1, 2), and for TiC14-Al(i-C4H9)3-piperylene for MMD narrowing (Fig. 5.13, curves 3,4).

129 M nlOd

0

i

0

20

40

60

100

80

w.mn

Figure 5.10. Dependence of M, on isoprene polymerization time T~ in the presence of TiCld-Al(iC4H&. 1 - free experiment, 2 - hydrodynamic effect on catalytic system in turbulent regime (variant 1, conditions as in Figure 5.7). MwlOd

2

0 .8

0,6

0A

--. 0

20

40

60

80

100

w.m n

Figure 5.11. Dependence of M w on isoprene polymerization time rP in the presence of TiCl4Al(i-C4H9)3-piperylene. 1 - free experiment, 2 - hydrodynamic effect on catalytic system in turbulent regime (variant 1, conditions as in Figure 5.7).

130 MnlO-5

I' -{ e

-. 0

20

40

60

100

80 @,Inn

Figure 5.12.Dependence of M, on isoprene polymerization time 2, in the presence of TiCLAl(iC4H9)3-piperylene. 1 - free experiment, 2 - hydrodynamic effect on catalytic system in turbulent regime (variant 1, conditions as in Figure 5.7). MwMn 11

-

9-

7 -

5

i

0

20

40

60

80

100 D,mn

Figure 5.13. Dependence of M,/M,, on isoprene polymerization time op in the presence of TiCl~Al(i-C4H9)3(1, 2 ) and TiC14-Al(i-C4H9)3-piperylene (3, 4). 1, 3 - free experiment, 2 , 4 hydrodynamic effect on catalytic system in turbulent regime (variant 1). Molecular masses increase under carrying out of stage of active sites formation of isoprene polymerization in the presence of Ti-AI catalytic system in turbulent regime may be explained after consideration of distribution of sites of macromolecules propagation by their kinetic activity (see 1.4.1). It is obvious from Figures 5.14 and 5.15 that hydrodynamic effect on catalytic system in

131

turbulent regime leads to redistribution of sites of macromolecules propagation with increase of activity of sites responsible for high-molecular polymer fractions formation. Reduction of "gelfraction" content in polymer under hydrodynamic effect on catalytic systems TiC14-Al(i-C4H9)3 from 2,8% to 1,5% mass and TiCl4-Al(i-C4Hg)3-piperylene from 1,4% to 1% mass (polymerization time T~ = 90 min) also testifies to redistribution of polymerization sites kinetic activity with the change of catalyst's particles surface structure.

2

1

1,6 1,2 0,8 0,4

0

Figure 5.14. Distribution of sites of macromolecules propagation by kinetic activity under isoprene polymerization in the presence of TiC14-Al(i-C4H9)3. 1 - free experiment; 2 hydrodynamic effect on catalytic system in turbulent regime (2, = 10 min).

1,6

1

9

2

11

l3

In(M) l 5

Figure 5.15. Distribution of sites of macromolecules propagation by kinetic activity under isoprene polymerization in the presence of TiCId-Al(i-C4H9)3-piperylene. 1 - free experiment; 2 - hydrodynamic effect on catalytic system in turbulent regime (2, = 10 min).

132

Thus, untraditional way of influence on process proceeding character, molecular characteristics and quality of resulted polymer product under isoprene polymerization initiated by Ti-AI Ziegler-Natta catalytic systems was revealed. Separation of fast stage of formation of macromolecules propagation sites and initiation with the use of tubular turbulent apparatus of jet type at significant increase of turbulence level at this stage and of slow stage of diene polymerization itself in tank reactors-polymerizers is advisable. In particular, preliminary turbulent mixing of reaction mixture components and Ziegler-Natta catalytic system under isoprene polymerization at the expense of use of tubular turbulent pre-reactor of divergent-convergent design allows changing of process rate in wide limits and determines the possibility of catalyst consumption reduction. Disperse structure of micro-heterogeneous Ziegler-Natta catalytic systems and consequently their activity are significantly effected by modification by diene additives that is essentially intensified by preliminary hydrodynamic effect on reaction mixture in turbulent regime. Moreover, stages separation allows significant improvement of molecular characteristics of resulted polymer products, changing of number and concentration of polymerization sites participating in polymer products synthesis. Obviously, separation of fast and slow stages with the use of tubular turbulent apparatus is necessary to realize also under polymerization of other dienes in the presence of micro-heterogeneous Ziegler-Natta catalytic systems. It was found that this method is applicable also for reception of ethylene-propylene rubbers, where the stage of solvent saturation by gaseous monomers proceeds in diffusion regime. 5.3.2. Ethylene and propylene copolymerization Works on investigation and optimization of physical-chemical processes of synthetic rubbers production are carried out for the long time. The bases for such investigations are the inconsistency between production volume and standard of existing equipment, the necessity of broadening of producing rubbers assortment and the most actual problem - technology imperfection and outmoded facilities of some production stages. From this point of view significant complications appear in particular under development of large-tonnage production of ethylene-propylene rubbers (SKEP and SKEPT) since the procrss is characterized by phase heterogeneity (presence of liquid and gas phases) and multicomponentness (ethylene, propylene, dicyclopentadiene (ethylidennorbornen), hydrogen, catalyst, cocatalyst, solvent). Ethylene-propylene rubbers reception process in acting productions is usually realized in reactorspolymerizers of volume about 16 m3 at intensive mechanical mixing. The height of reaction volume infill is 60%. Introduction of reaction mixture components directly into stirred reactor of large volume with mixer (Fig. 5.16) as a rule does not provide uniform saturation of liquid products by monomers and hydrogen that due to diffusion limitations appearance leads to broadening of MMD of resulted polymer products (see 1.4.3). It is known that in any heterophase process reagents interaction takes place at phases interface or in immediate vicinity to it. That is why in particular under ethylene-propylene rubbers synthesis for increase of effectiveness of liquid products saturation by monomers and hydrogen maximum increase of phase contact surface is required. In general case it can be reached by increase of dispersed system elements (drops, bubbles) number and decrease of their sizes at given volume of dispersed phase. So, application of compact tubular turbulent apparatus of jet type of divergent-convergent design (Fig. 3.1b)is advisable under ethylene-propylene rubbers production. In particular, tubular turbulent apparatus of divergent-convergent design may be used at the stages of homogeneous gas-liquid mixture preparation with its following introduction into parallel functioning tank reactorspolymerizers (Fig. 5.16). Thus, in this case we suggest to use the method of separation of fast and slow stages considered by the example stereoregular cis-l,4-polyisoprene synthesis. However, in the case of ethylene-propylene rubbers the fast stage is mass-exchange physical process, that is solvent saturation by gaseous monomers. With the aim of estimation of effectiveness of use of tubular turbulent apparatus of divergentconvergent design for preparation of homogeneous gas-liquid mixture with developed surface of

133

phase contact under ethylene-propylene rubbers synthesis it is advisable to study the quality of resulted disperse systems in apparatus of this type (Fig. 3.lb) in comparison with stirred tank reactors (Fig. 5.16). Undissolved part, gas

gas (monomers mixture)

Figure 5.16. Stirred tank reactor for synthesis of ethylene-propylene rubbers. 1 06beMHbIfi PeaKTOp CMeLUeHMIl AnIl CHHTe3a 3TPUIeHllpOnHneHOBbIX KBY'IYKOB. 1 - heat-carrying agent jacket; 2 - mechanical mixer; 3 -bubbler. On the base of experimental data the equation for calculation of average diameter dz of disperse phase particles in mixing zone for "liquid-gas'' system was derived for stirred tank reactor [ 1161: Q6 2

(5.5)

0

Here N - mixer power spent on liquid mixing, n - number of mixer revolutions, d M - mixer's diameter, Gz - gas volume throughput, w - volume of gas-liquid mixture in apparatus. Equation of (5.5) for sure totally reflects the character of influence of mixing intensity on sizes of dispersion in system "gas-liquid". By order of value the calculated values of dz are comparable with those calculated by empiric equation [ 1291: -314 112

2 d = ( 2 . 3 +~ 0.75)

.[ !k1°'4

(5.8)

dM

where D - tank reactor volume, Rec'= plndizIp - centrifugal Reynolds' criterion, Wec = n2d,pl/o centrifugal Veber's criterion.

134

At values of continuum phase viscosity, in our case of liquid p1= 0,001+0,0025 Pasec the average gas content can be calculated by equation [ 1161: - 0,3

cp = 0,33

.1[ '

0,s

.[ 2) .[E] 0,17

(5.9) v"s For calculation was accepted that V,, = 0,265 d s e c . The average diameter of disperse phase in mixer rotation zone calculated for system "liquidgas" by (5.5) and (5.8) was 1,2 mm. However outside the mixing zone the averaged diameter of bubbles is quickly increased due to the reduction of turbulence level and consequently of intensive bubbles coalescence. In peripheral part of tank apparatus the average diameter of disperse phase particles dz in system "liquid-gas'' can be calculated by empirical dependence [116]: (5.10) .

.

In this region of stirred tank reactor the average bubbles diameter significantly increases and is equal to 8 mm. Calculations on the base of equations received in present work (3.40)-(3.43) (see 3.2.1) showed that size of disperse phase particles averaged by tubular turbulent apparatus volume for concrete conditions of industrial production of SKEP(T) was equal to 0,127 mm, whereas using standard stirred tank reactors (Fig. 5.16) under barbotage it is in order higher - 1,2 mm. Thus, application of compact tubular turbulent apparatus of divergent-convergent design at stage of formation of homogeneous gas-liquid mixture before its introduction into tank polymerizer leads to the significant (practically in order) increase of phase contact surface. Developed phase interface leads to uniform saturation of liquid products by monomers and hydrogen that provides in this case improving of operational characteristics of received ethylene-propylene rubbers contrary to stirred tank reactors use. Spectra of pressure relaxation (SVPR) of SKEPT with DCPD samples melts obtained at standard conditions and under preparation of gas-liquid mixture in tubular turbulent apparatus of divergent-convergent design are presented in Figure 5.17. It is obvious that SVPR of copolymer obtained under tubular turbulent apparatus use has good bell-shaped form. This testifies to rubber micro-uniformity .

135

i

,MPa

-4

-2

0

2

4

6 lm, nin

Figure 5.17. Spectra of pressure relaxation time SKEPTJO DCPD. 1- stabdard scheme; 2 preparation of gas-liquid mixture in tubular turbulent apparatus. Homogenization of entering the polymerizer flows leads to reduction of blockness degree of ethylene units at the expense of long blocks content reduction (in 2,5 times) at maintenance of number of short units; to the increase of blockness degree of propylene units including increase of long blocks length, to the reduction of branching degree of macromolecules nearly in 2 times (Table 5.4).

Table 5.4. The influence of turbulent mixing of gas-liquid flows on SKEPT (ENB) macromolecules structure

Moreover, stage-by-stage preparation of gas-liquid mixture in tubular turbulent apparatus of divergent-convergentdesign and distribution of its introduction into parallel functioning polymerizers in tubular turbulent apparatus-dispenser of "spider" type allows reception of copolymers with equal properties in various parallel functioning polymerizers (Table 5.5).

136

Way of introduction of gas-liquid mixture Separate introduction Introduction by tubular turbulent apparatus-distributor

Parallel working polymerizers I I1 I I1

Mooney viscosity (48-50)+2 (46-47)rtl (50-52)+1 (50-52)+1

Thus, increase of turbulence level at stage of mixing of gas (ethylene, propylene, hydrogen, circulating gas) and liquid (solvent, DCPD or ENB) products in production of ethylene-propylene rubbers (SKEP and ENB) (separation of fast and slow stages) determines guaranteed possibility of reception of copolymer homogeneous in structure. Results received in 3.2 on intensification of turbulent mixing in multi-phase flows allow calculation of tubular turbulent apparatus construction and selection of optimal regime of their work in relation to ethylene-propylene rubbers production. In chemical technology processes of suspension reception by condensation method under proceeding of chemical reactions with precipitation are prevalent. Among polymer synthesis such processes as preparation of Ziegler-Natta micro-heterogeneous catalytic systems, anti-agglomerates for synthetic rubbers (on the base of metals carboxylates, in particular calcium stearate), etc. are related to them. Literature data review shows that at present there are no works concerning the possibility of use of tubular turbulent apparatus for realization of fast chemical reactions with combination of homogeneous liquid phase flows and suspensions formation. As a consequence, it seemed to be advisable to study the possibility of tubular turbulent apparatus of divergentconvergent design use for receipt of homogeneous fine suspensions by condensation method under fast chemical reaction proceeding and to compare the effectiveness of their work with tank apparatus with mechanical mixing.

5.4. Fast chemical reactions followed by solid phase formation The method of suspension receipt by mixing of initial homogeneous reagents in turbulent regime realized by rotary pumps arms is prevalent at conditions of industrial production for carrying out of fast chemical reactions. With the aim of comparison of effectiveness of work of tubular turbulent apparatus and tank reactors (rotary pumps) mechanical mixer of "turnip" type was chosen that worked by rotary pump principle. In combination with intensive circulating movement in vertical direction (Fig. 5.18)

Figure 5.18. Vector field of rates in reaction volume under work of high-speed mixer of "turnip" type (- 200 revs).

137

Microscopic analysis of received suspensions showed that particles of Bas04 posses pronounced anisometry (they represent sticks with length-to-diameter ratio about 4, that explains the sigmoid shape of sedimentation curves with extreme) (Fig. 5.19). At the initial moment of sedimentation the rotation of stick-shaped particles is possible, that determines additional resistance to sedimentation proceeding (analogously to viscosity increase) and slowing down of rate of deposition accumulation. Moreover, at free precipitation particles of spherical form are oriented in movement direction in that way to create maximum resistance to movement. This also reduces the precipitation rate of solid particles in liquid and embarrasses the definition of their actual sizes. In this connection, equivalent radius re (radius of spherical particle precipitating with the same rate) was determined by results of sedimentation analysis. It is obvious from the shape of sedimentation curves (Fig. 5.19) that at rise of mixing rate they shift to the region of particles with smaller sizes (higher sedimentation times), i.e. under increase of turbulence level of mixing decrase of average particles radius in reaction zone occurs.

mI-% 300

200

100

0 0

100

200

300 7, I1-dn

Figure 5.19. Curves of sedimentation of Bas04 particles. Mixing rate (variant 2, see 2.2.6): 300 (1); 1500 (2); 2500 (3) revs; 4 - tubular turbulent apparatus. Points - experiment; lines approximatingfunction (2.3). It is obvious (Fig. 5.20) that in the case of application of stirred tank reactors for suspensions receipt (variant 1 and 2, see 2.2.6) there is the reduction of the most probable particles radius rmpat the rise of mixing rate N (from rmp= 4,6 mcm in the absence of mixing down to rmp= 1,2 mcm at N = 500 revs). In this case reaction mixture according with experimental conditions is diluted enough and increase of turbulent mixing can not significantly increase the rate of crystal growth and consequently their sizes. Significant reduction of disperse phase particles radius occurs at the expense of high rate of micro-mixing intensity of which is increased at the rise of dissipation of specific turbulence energy density dissipation E at increase of mixing rate (see 3.1.4.2). At high rate of micro-mixing nucleation rate-to-crystal growth rate ratio is increased that leads to the rise of nucleuses number of new phase and decrease of rmp.This fact is in particular observed in the case of use of baffle wall (variant 2, see 2.2.6) increasing the intensity of turbulent mixing in tank apparatus. In this case the relative rate of reduction of rmpwith increase of mixer revolutions number N is higher in comparison with the first synthesis method (Fig. 5.20, curve 2). It is noticeable that rmpwith increase of mixing rate in tank reactor (flask) up to the maximum possible in laboratory conditions (N = 5000 revs) as in both variant 1 and 2, approaches to size of suspension

138

particles rW = 1,l mcm received in tubular turbulent apparatus where the high level of turbulent mixing is created only at the expense of canal geometry and flow hydrodynamic energy (see Chapter 3).

0

I

I

1

m

4000

m N,IWO

Figure 5.20. Dependence of the most probable radius rnlpof Bas04 particles on mixing rate. Tubular turbulent apparatus (l),variant 1 (3), variant 2 (2, 4); after preliminary formation in the absence of mixing and aging for 90 min (4). With the aim of estimation of polydispersity degree of resulted suspensions by sedimentation curve (Fig. 5.19) weight-average rwand numerical-average r, particles radiuses were calculated: (5.11) where nwi,n~ - weight and numerical part of particles with radius ri. for homogeneous suspension polydispersity coefficient was rwI r, = 1. Under increase of mixing intensity in the course of solid particles formation in system according with variant 1 (see 2.2.6) there is the decrease of polydispersity coefficient from rwI r, = 2,8 in the absence of mixing down to rwI r, = 1,55 at N = 5000 revs (Fig. 5.21). Homogeneity of suspension received according with variant 2 is practically doesn't change with the increase of turbulent mixing intensity and is equal to rw/ rn= 1,77. Carrying out of fast reaction with precipitation in tubular turbulent apparatus allows formation of suspension with homogeneity high enough rwI rn = 1,48 (Fig. 5.21). Naturally, in the course of work it seemed to be advisable to study characteristics of suspensions effecting by hydrodynamic action not at the moment of their formation, but after their preliminary aging in the absence of mixing (modelling of hydrodynamic action on Ti-AI catalytic systems in turbulent regime (see 5.3.1)).Interaction of initial components (water solutions of HzS04 and BaC12) with this aim was carried out in the absence of mixing with aging for 90 min and further influencing on received suspensions of turbulent mixing according with variant 2 with the use of tubular turbulent apparatus. It was found that in this case average radius of particles doesn't increase while time of aging is increased (Fig. 5.20, curve 4). At the increase of intensity of turbulent mixing of mixture containing preliminary formed suspension sharply reduction of radius of solid particles received in the absence of mixing (Fig. 5.20, curve 4)doesn't occur, in comparison with suspension formed at various mixer rotation rate (Fig. 5.20, curve 2, 3). however, fine particles with rnlp= 1,2 mcm are firmed at intensive hydrodynamic action in tank apparatus (N = 5000 revs). The use of tubular turbulent apparatus for effecting on aged for 90 min suspension allows reception of disperse

139

systems with less particles sizes (r,,,,, = 1 , l mcm) (Fig. 5.20, curve 1). Analogous picture is observed also for polydispersity coefficient rw I rn of resulted suspensions (Fig. 5.21). The use of tubular turbulent apparatus both at the moment of suspension synthesis and under the influence on preliminary formed and aged for 90 min disperse system leads to receipt of particles homogeneous enough with polydispersity coefficient rw/ rn= 1,48 (Fig. 5.21, curve 1).

2000

6000

4000 N,EVS

Figure 5.21. Dependence of polydispersity coefficient rw/ rn of Bas04 particles on mixing rate. Tubular turbulent apparatus (1); variant 1 (3); variant 2 (2, 4); after preliminary formation in the absence of mixing and aging for 90 min (4). Thus, use of tubular turbulent apparatus of divergent-convergent design both at synthesis moment and treatment of ready suspension of Bas04 allows reception of fine systems with average radius of disperse phase particles rnlp= 1,l mcm and polydispersity coefficient rwI rn = 1,48, that is reached in tank apparatus with mechanical mixer only at mixing rate higher than 2000 revs. Advisability of compact high-productive tubular turbulent apparatus use for realization of fast chemical reactions with receiving of homogeneous fine suspensions was revealed. In this case time of reaction mixture residence in apparatus is about 2-3 sec that reduces possibility of secondary processes proceeding. These apparatus are effective not only at synthesis but also at treatment of preliminary formed suspensions. As a consequence, it is necessary to use tubular turbulent apparatus for receipt of Ti-AI, V-AI, Nd-AI and other high-active micro-heterogeneous Ziegler-Natta catalytic systems in processes of (co)polymerization of olefins and dienes (cis- 1,Cisoprene, ethylenepropylene, cis-1,Cbutadiene and other rubbers) (that is experimentally confirmed, see 5.3.l), antiagglomerates (stearates of Me2+)for synthetic rubbers, pigments, in particular on the base of Ti02 and other homogeneous fine suspensions resulting from the interaction of homogeneous liquid flows. Obviously apparatus of this type is not inferior in effectiveness to pump or other power-inputting equipment for suspension synthesis. Obtained results open wide possibilities for modification of heterogeneous Ziegler-Natta catalytic systems at the expense of hydrodynamic action on disperse structure of catalyst particles. The results obtained in this chapter significantly deepen the novel part of chemical physics concerning fast chemical processes proceeding in turbulent regime in particular under polymer synthesis. We showed that increase of level of flows turbulent mixing significantly influences on character of chemical process proceeding and quality of synthesized products (polymers and

140

copolymers structures, their molecular characteristics, disperse structure of catalytic systems, etc.). Results obtained in this work allow development of new solutions in the region of fast chemical processes realization. It was found that conception of use of compact, high-productive, energy- and resource-saving tubular turbulent apparatus is applicable not only for novel class of chemical processes - fast liquid-phase reactions proceeding in diffusion regime. Differences in kinetic parameters under proceeding of one and the same reaction in the presence of catalysts differing in activity determine the necessity of various approach to the selection of type and construction of main reactor. In this case the preference should be given to the most active catalytic system that allows carrying out of the process in turbulent flows limited by impenetrable wall. Possibility of realization in turbulent regime of both fast polymerization processes and fast polymer-analogous reactions under polymers modification was revealed and confirmed by the example of butyl rubber chlorination in solution by molecular chlorine. Interesting and practically important idea of separation of olefine and dienes (co)polymerization processes initiated by Ziegler-Natta catalysts to fast (active sites and reaction mixture formation) and slow ((co)polymerization itself,) stages was proposed. In practical aspect this fact determines the advisability of use of tubular turbulent apparatus before stirred tank reactor with high residence time for realization of the main process - polymerization. We showed that realization of fast chemical reactions with solid phase formation is the new and perspective field for tubular turbulent apparatus application. This opens wide possibilities for Ziegler-Natta catalytic systems modification at the expense of influencing on disperse structure of micro-heterogeneous catalyst' particles both at synthesis moment and under treatment of preliminary aged complex. The obtained fundamental regularities of fast processes realization in turbulent regime that are presented in Chapters 3-5 serve as a base for creation of a number of energy- and resource-saving high-productive technologies on the base of tubular turbulent apparatus.

141

CHAPTER 6. DEVELOPMENT OF NOVEL TECHNOLOGIES FOR FAST PROCESSES IN TURBULENT REGIME 6.1. Plug-flow tubular turbulent apparatus - the novel type of industrial apparatus Novel apparatus types for effective fast chemical reactions realization in turbulent regime were worked out (Fig. 6.1, Table 6.1). For liquid-phase processes methods for calculation and construction of high-productive compact tubular turbulent apparatus of plug-flow of jet type but working in turbulent flows were developed. Apparatus combine the best technical and technological advantages that are known in industry of mixing and plug-flows reactors and are differ by the presence of their own original inhering only to them fundamental particularities. This type of apparatus is not considered in active classification of chemical technology apparatus. Four modifications of tubular turbulent apparatus irreplaceable under realization of fast chemical processes and suitable for carrying out of mass-exchange liquid-phase processes in industry were developed: cylindrical (with constant reaction zone diameter along apparatus length); shell-and-tube (pipe bundle washing by coolant); "zone" (concatenate independent adiabatic reaction zones separated by cooling zones) and divergent-convergent (with several concatenate sections with local hydrodynamic resistances) designs principally differing from known industrial apparatus of mixing and plug-flow, by this tubular turbulent apparatus may function in both regimes plug-flow and mixing.

1 reagent 2

coolant

J p -= -~ ~ o d u 1 1 reagent2

+1 reagent

3

"

+

- J-

coolant I reagent 2

reagend+l/

- v

+

v A

reagent 2

4

--

b,

products

c)

Figure 6.1. Tubular turbulent apparatus of cylindrical (a), divergent-convergent (b), shell-and-tube (c) and "zone" (d) construction (1 - reaction zone, 2 - cooling zone). Relative disadvantages of compact tubular turbulent apparatus of the first generation are: fast decay of turbulent diffusion coefficient Dt in reaction zone and along its length (over 1-2 calibres) (see Fig. 3.6a) and low effectiveness of external heat removal (see Fig. 4.20). In a number of cases they succeeded in intensification of external heat removal. Shell-and-tube turbulent reactors were

142 142

worked out with this aim. Using pipe bundle washing by coolant and maintaining required productivity one may significantly improve external heat removal especially under process carrying out in quasi-isothermal conditions. If under realization of fast chemical processes on the base of two constructions of turbulent apparatus effective heat removal can not be provided then for the avoidance of inadmissible overheating of reaction mixture one may use "zone" model. The reagents (or catalyst) portion that determines admissible temperature rise in reaction zone is introduced into every zone, where reaction proceeds. The bigger the number of reaction zones, the bigger the possibilities of fast chemical reactions control. Table 6.1. Indexes* of reactors work of continuous action under fast chemical reactions

catalyst water raw materials

1 038 0,75

1 0,8

0,75

1,5-2 1 1

1,3 079 0,85

Tubular turbulent apparatus of the second order of divergent-convergent design that essentially by all positions surpass any known apparatus of chemical technology are the best in their technical characteristics. In particular, their distinctive features are: possibility of receipt of required turbulence level and consequently intensity of mixing and heat-transfer, maintaining of practically constant turbulence level in apparatus volume and along its length at significantly lower flow rate (Fig. 3.6b); possibility of formation of self-similar regime in relation to viscosity and criterion Re (see Fig. 3.18), etc.. As a consequence apparatus of this type are irreplaceable for work with any mediums including high-viscosity polymer solutions and allow increasing of process productivity; high effectiveness under work with multi-phase systems because there is fine phases dispersion with receipt of homogeneous emulsions and foams due to geometry of divergent-convergent canal, significantly higher (in order and higher) solution rate and gas amount in liquid in comparison with barbotage or intensive mechanical mixing; increase (in 2-4 times) of coefficient of heat-transfer

143

through the outside wall in comparison with apparatus of cylindrical design. Fast process proceed in optimal regime if the ratio Tch > Tturb > rmeso > rmicro is met, by this decrease of reaction zone radius R leads to reduction of characteristic mixing times that represents the key to realization of fast processes in optimal conditions. Change of liquid flows rate in apparatus of divergent-convergent design and ratio dd / d, is the only but effective way of influence on the dispersion character and quality of resulted emulsions. Decrease of dd and increase of turbulence level in reaction volume (at the expense of V, Dt and E increase) promote receipt of fine systems. Apparatus of divergentconvergent design may function in both plug-flow and mixing regimes that is reached firstly by the change of apparatus geometry and reaction mixture movement rate. Formation of circulation zones provides flows in apparatus of this type. Under movement of multi-phase mediums under the action of centrifugal forces the separation effect may be observed leading to the appearance of heterogeneities in components distribution in disperse system with phase separation in dependence of density of participating in process reagents. Thus, tubular turbulent apparatus of any construction under formation in them of quasi-plugflow mode in turbulent flows represent principally novel type of industrial compact high-productive apparatus combining the best technical and technological advantages of jet stirred and plug-flow reactors using in industry at present and having their own original particularities. Realization of tubular turbulent apparatus under fast processes carrying out at conditions of industrial production allows creation of continuous energy- and resource-saving rechnologies.

6.2. Polymers production 6.2.I . Ethylene-propylene rubbers production On the base of results received in this work the scheme of continuous solution copolymerization of ethylene and propylene (Fig. 6.2) was developed and introduced into production (OAO "Nizhnekamskneftekhim"), tubular turbulent apparatus are used in it at the following stages: 1) preparation of homogeneous gas-liquid mixture and its introduction into parallel working polymerizers (Fig. 6.2, positions 2-4, Fig. 6.3, positions 8,9, 10, 12); 2) formation of sites of macromolecules propagation (Fig. 6.2, position 6). (Patent of Russia No.2174128; priority of 27.03.2000, published on 27.09.2001, bulletin 27).

144

Figure 6.2. Principle scheme of SKEP(T) production. 1 - comb of preliminary mixing; 2, 3, 4 , 6 , 13, 14 - tubular turbulent apparatus; 5, 9 - polymerizers; 7, 10 - condenser of blowing back; 8, 11 - collector condenser blowing backs; 12 - separator-eliminator of liquid phase of blowing backs.

Figure 6.3. The method of continuous solution copolymerization and reactor-distributor for it realization. 1 - polymerizer; 2 - drive of mixing device; 3 - connection pipe for introduction of gas-liquid mixture; 4, 5 - connection pipes for separate introduction of catalytic complex components; 6 - connection pipe for copolymer taking off; 7 - comb; 8, 9, 12 - tubular turbulent apparatus of divergent-convergent design; 10 - tubular turbulent apparatus-distributor; 11 - transfer tubes; 3 - connection pipes for introduction of gas-liquid mixture into polymerizer.

145

Under formation of homogeneous gas-liquid mixture (Fig. 6.3) cooled components in definite proportions (see 2.2.4) under pressure are introduced into comb (position 7) and tubular turbulent apparatus of divergent-convergent design (positions 8 and 9) in which developed turbulent regime of flow and consequently preliminary mixing are provided. Gas-liquid mixture again mixers at conditions of turbulent flow when getting into tubular turbulent reactor-distributor (position 10). Gasliquid mixture is introduced into transfer tubes (position 11) practically does not change conditions of movement. Then, via tubular turbulent apparatus of divergent-convergent design (position 12) gas-liquid mixture is introduced into parallel working polymerizers. Developed turbulent flow at preliminary stage provides uniform distribution of components in gas-liquid mixture (position 8 and 9). Tubular turbulent apparatus-distributor (position 10) provides maintaining of homogeneity of gas-liquid mixture components in volume at their distribution among transfer tubes (position 11). This forms equal structure of gas-liquid mixture introducing into parallel working reactors-polymerizers. Gas-liquid mixture before introduction into polymerizer is again subjected to turbulization for providing homogeneity of components distribution. Stage-by-stage preparation of gas-liquid mixture in tubular turbulent apparatus and distribution of its introduction to the parallel working polymerizers allows receipt of copolymers with equal properties in various reactors and at other equal conditions of polymerization the increase of total process productivity was observed and productivities of separate polymerizes by this became identical (see 5.3.2). At the stage of formation of sites of macromolecules propagation under ethylene with propylene copolymerization initiated by V-A1 catalytic systems (separation of fast stage of active sites formation and initiation and slow stage of copolymerization itself) in novel technological scheme of SKEPT production tubular turbulent apparatus is installed (Fig. 6.2, position 6, Fig. 6.4). Calculation of geometry of zone of catalytic system components mixing is also carried out on the base of quantitative dependences obtained in Chapter 3.

Figure 6.4. Tubular turbulent apparatus for formation of homogeneous high-dispersed sites of macromolecules propagation under ethylene and propylene copolymerization.

In this case catalyst is introduced via nipple 1 by inside canal 2 of pre-reactor (Fig. 6.4) and solvent via nipple 3. The flow after passing via turbulization section 4 equalizes catalyst concentration over flow volume and introduces into 5. By inside canal the flow comes to the divergentconvergent section 6 where its turbulization occurs and then to the nozzle 7, where flow additionally acquires non-stable hydrodynamic current. At output of nozzle the flow reaches Veber's criterion 0,5 and as a result of interaction with gas medium begins to break but not earlier than at 50 mm from nozzle head. Cocatalyst is introduced via socket 8 into outside canal 9 in which via socket 10 the solvent is preliminary introduced. The flow passing turbulization section 1 1 equalizes cocatalyst concentration over volume and is introduced into outside canal 5. In outside canal the flow is also preliminary turbulized in divergent-convergent section 12 and then via nozzles 13 is introduced into gas me-

146

dium. Getting out from head of nozzle 13 the flow has value of Veber's criterion 0,5 and as a result of interaction with gas medium is broken. Breaking process should start not earlier than at 50 mm from nozzle head. Fast dilution of catalytic complex components is provided by turbulization of mixing flows that is realized by configuration of inside surface of each canal of tubular turbulent pre-reactor. By this the flow by turns undergoes pressing and broadening promoting formation of turbulences that provide full and fast mixing of liquids with various density over the whole volume of moving flow. Mixing efficiency is provided in at least two divergent-convergent sections. Under catalytic complex components mixing fast reaction of V-AI catalyst formation proceeds and growing crystal active sites of high activity is as a rule precipitated onto metallic surface of device applying for components mixing. Gradually the thickness of deposited layer is increased, copolymerization reaction is additionally proceeded on it and polymer layer is formed that leads to driving of holes of catalytic complex solutions introduction. Moreover, formed at these conditions catalytic complex due to the absence of steric difficulties forms large crystals with perfect structure covering the most part of active sites inside of the crystal. That is why creation of conditions of separate introduction of catalyst components in tubular turbulent pre-reactor of developed construction (Fig. 6.4) excluding the deposition of crystals on apparatus surface. Moreover, creation of conditions of flows interaction mainly in gas phase over reaction mass surface on the one hand creates conditions of catalytic complex components fast mixing with formation of small catalyst particles, on the other hand - formation of these crystals occurs in gas medium containing large amount of gaseous monomers which are precipitated on active sites of growing crystal faces preventing from formation of perfect and large crystals. Results of carrying out of stage of V-A1 catalytic system active sites formation in turbulent regime are presented in Table 6.2. Table 6.2. Formation of active sited in turbulent regime under ethylene and propylene copolymerization initi-

SKEPT with DCP

SKEP

24 335 750 43-45k1 48 3,s 780 45-47+1 1-5 - formation of active sites in turbulent regime with use of tubular turbulent pre-reactor (Fig. 6.4) 6-7 - separate introduction of catalytic system components directly into tank reactor-polymerizer. Preparation and introduction into polymerizer of catalytic complex components in tubular turbulent pre-reactor (Fig. 6.4) provides production of copolymer (SKEP and SKEPT) with good and stable in time properties (Table 6.2). Analogously to laboratory results received under separation of stage of formation of active sites and initiation of isoprene polymerization on Ti-AI catalytic system (see 5.3.1) at conditions of industrial production under ethylene of propylene copolymerization increase of copolymer yield and its molecular mass and also decrease of catalyst consumption are observed.

147

6.2.2. Chlorobutyl rubber production On the base of studying of specificity of process of butyl rubber chlorination in solution by molecular chlorine, conditions of formation of self-similar field of liquid flows and regularities of multi-phase flows in tubular canals (see Chapter 3) compact tubular turbulent pre-reactor-chlorator (Fig. 6.5) with original nozzles (Fig. 6.6) was developed and approved on OAO "Nizhnekamsneftekhim" (Patent Russia No. 2170237, priority from 30.07.1999, published on 10.07.2001, bulletin No. 19; Patent Russia No. 2170238, priority from 30.07.1999, published on 10.07.2001, bulletin No. 19). Nozzles are placed in series along tubular reactor and have on the outside surface the flutes forming with inside surface of box wall screw canal (Fig. 6.6), and inside surface of nozzles has on flanks conical lugs with diameter in 1,3-3 times lower than in cylindrical part of inside cavity of nozzles and forming tubular turbulent reactor of divergent-convergent design (Fig. 6.5). Screw canal and inside cavity of nozzle are connected by holes and/or cuts, and beyond last nozzle in the course of gas-liquid flow one or more static mixers are placed. Forming in tubular turbulent reactor of divergent-convergent design gas-liquid mixture is separated in placed in series along the flow nozzles into inter-communicating flows that are united into common flow at the output of last nozzle. Taking into account that the volume of gas medium in process exceed the volume of BR solution after preliminary mixing of gas-liquid mixture one should separate the flow into at least two parts - it is divided in the nozzle into peripheral and axial flows that provide creation of optimal conditions for chlorination reactions proceeding.

Figure 6.5. General form of tubular turbulent reactor of divergent-convergent design with nozzles for chlorination of BR. 1 - apparatus box; 2 - connecting pipe for introduction of butyl rubber solution; 3 - connecting pipe of gas mixture or gas introduction; 4 - flow distributor; 5 - nozzles; 6 - static mixer; 7 - connecting pipe for gas-liquid mixture taking off.

148

In production of chlorobutyl rubber it is advisable to use tubular turbulent apparatus of cylindrical or divergent-convergent designs also at other stages of technological process, in particular under neutralization of resulted chlorobutyl rubber solution, scouring of chlorobutyl rubber solution by water from salts and other compounds, washing of inverse solvent; under introduction into chlorobutyl rubber solution of stabilizer and anti-agglomerator and also instead of all intensive including involume agitators with mechanical mixers (Fig. 6.7).

Figure 6.6. The scheme of nozzle in tubular turbulent reactor. 1 - hollow box; 2 - screw thread; 3 - inside cavity; 4 - conic lugs (convergents); 5 - holes; 6 - cuts. Thus, principally novel, unique, economic, continuous process of chlorobutyl rubber production with the use of tubular turbulent apparatus of original construction operating in regime of high flows turbulence on at least four stages of technological scheme was worked out (Fig. 6.7).In novel process stirred tank reactors are substituted by tubular reactors at stages of butyl rubber chlorination (position 4), neutralization (position 5), stabilizer and anti-agglomerator introduction (position 12, 15). It is possible to replace by tubular apparatus the flush column also (position 9) where water washing of solvent is carried out. The process is remarkable for compactness and guarantees significant energy- and resource-saving, high environmental safety, easy control, process productivity increase (in hundreds of times), significant decrease of production space, increase of products quality. Technology of chlorobutyl rubber production is significantly simplified, investment expenses and products prime cost are decreased. Process is continuous. Novel process on the base of compact tubular reactors is principally fitted for production of bromobutyl rubber and also others chlorinated elastomers. Process is accepted at OAO "Nizhnekamskneftekhim". In 200 1-2002 pilot plant worked at which chlorobutyl rubber of required quality was received without any problems. Technical schedule was made. At present experimentalindustrial plant is projected.

149

.,

111

Ill

I

4

5

1

w

t

I

I

24

++ XI11

XIV

+

VI1

xv

Figure 6.7. Technological scheme of chlorobutyl rubber production with the use of tubular turbulent apparatus ofjet type: 1 - apparatus of chlorobutyl rubber solution preparation; 2, 8, 11, 14, 19, 21 - pumps; 12, 15 - tubular turbulent apparatus for introduction into chlorobutyl rubber of stabilizer-antioxidant and anti-agglomerator; 4 -tubular turbulent apparatus-chlorator; 5 - tubular apparatus-neutralizer; 6 - filter; 7 - collector; 9 -tubular turbulent apparatus for water washing of solvent; 10, 24 - sediment boxes; 13 - neutralizer; 17 - degasifier; 18 - orifice device; 20 - vacuum degasifier; 22, 23 - condenser. Flows: I - solvent; I1 - butyl rubber crumb; I11 - chlorine; IV - nitrogen; V - alkali water solution; VI - water; VI1 - washing water for purification; VI11 - stabilizer-antioxidant solution; IX - anti-agglomerator suspension; X - vapor; XI - coolant; XI1 - to the vacuum line; XI11 - benzene dewatering; XIV - water for stripping of organic compounds; XV pulp into concentrator.

6.2.3. Liquid oligo-piperylene rubber production Traditional technological scheme of liquid oligo-piperylene rubber of SKOP sort production [ 176, 1771 independently on catalytic system activity includes as main apparatus stirred tank reactor (1 2 m3) or cascade of them (in 1,5 m3) equipped by mechanical mixers and also by outside and inside coolers (coolant is inverse water) (Fig. 6.8a).

150

w

V

+

1

1

&I

5

I

t

t IV b)

Figure 6.8. The scheme of liquid oligo-piperylene rubber of SKOP sort synthesis. 1, 2 - tank reactors-polymerizers with frame mixer and cooling devices; 3 - tank apparatus-degasifier (a); 1, 3 compact tubular turbulent apparatus for oligo-piperylene synthesis and catalyst deactivation; 2 cooler (b), 4 - sediment box, 5 - degasifier. I - solvent, I1 - catalyst, I11 - piperylene fraction, IV deactivator, V - recycle, VI - SKOP for storage. Kinetic analysis of main reactor in dependence on used catalytic system showed (see 5.2) that under oligomerization of piperylene initiated by AIC13.0(CsH5)2 it is advisable to use tubular turbulent apparatus. Not high values of dynamic viscosity of liquid oligo-piperylene rubber at output of reactor (at 283 K, p = 1 mPa.sec) allows also perfecting of stage of catalyst deactivation by propylene oxide or water (Fig. 6.8b). Novel technology of piperylene oligomerization (Fig. 6.8b) guarantees saving of main principles of working technological schemes, increase of specific productivity of main reactor, reduce of specific weight of side reactions and secondary process proceeding at the expense of significant reduce of zr (Table 5.3). On the base of received in this work results recommendations on improving of piperylene oligomerization with the use of tubular turbulent apparatus were given for A 0 "Kauchuk" (Sterlitamak city) and OAO "Nizhnekamskneftekhim".

151

6.2.4. Antiagglomeratorproductionfor synthetic rubbers At present calcium stearate (anti-agglomerator) is produced by two-stage method: C17H35COOH + KOH -+ C17H35COOK + H2O 2C17H35COOK+ CaC12 -+ (C17H35C00)2Ca + 2KC1 The second stage - interaction of potassium stearate (water solution) with calcium chloride (water solution) is the reaction of ion exchange and proceeds with high speed, i.e. in diffusion regime that required intensification of mass-exchange process. Received in this work results on fine suspension synthesis (see 5.4) allow recommending of application at this stage the tubular turbulent apparatus of divergent-convergent design (Fig. 6.9) consisting of cascade of three concatenated mini-reactors of jet-type. Proposed construction of apparatus is characterized by high value of turbulence kinetic energy density dissipation and consequently by minimum micro-mixing characteristic time (see Chapter 3). This determines the possibility of creation of local shifting deformations and receipt of fine suspensions (see 5.4).

R,f 35Nl i E

.
Ls

\

A I

- 4

L apparatus 1

L apparatus 2

L -1

apparatus 3

Figure 6.9. General view of tubular turbulent apparatus for stage of interaction of potassium stearate with calcium chloride. Apparatus 1: dd = 30 mm, d, = 20 mm, L, = 90 mm, L = 0,9 m, y = 35+5', sections number 10; apparatus 2: dd = 40 mm, d, = 20 mm, L, = 120 mm, L = 1,4 m, y = 35+5', sections number 12; apparatus 3: dd = 50 mm, d, = 25 mm, L, = 150 mm, L = 3 m, y = 35+5', sections number 20. Total apparatus length is 5,5 m. Every section of turbulent reactor is calculated upon optimal operation at various process productivity: apparatus 1 - up to 2,5 m3/hour of summary liquid flows; apparatus 2 - 2,5-5 m3/hour; apparatus 3 - 5- 10 m3/hour. Proposed construction provides anti-agglomerator synthesis carrying out in a wide diapason of process productivity 2,5-10 m3/hour. Application of novel technology at stage of interaction of potassium stearate with calcium chloride allow expecting the following: stoichiometric reaction proceeding, decrease of reaction mixture residence time in apparatus, receipt of suspension of calcium stearate of high quality (firstly by dispersion), simplicity of reactor exploitation, increase of apparatus exploitation time between shutoffs due to drivings by reaction products. Corresponding recommendations were given for production of cis-l,4-isoprene rubbers (workshop I-5V, A 0 "Rubber", Sterlitamak city). 6.3. Low-molecular products production in the flow 6.3.1. Unblended benzene production

It was shown in chapter 3 that flows of liquid components differ in density and viscosity at least in 20%, in particular infinitely mixing glycerin (p = 1,26 g/cm3) and water (p = 1 ,O g/cm3), are segregated in tank apparatus. So, creation of homogeneous phase is a problem. On the plant "Ethylene" of OAO "Nizhnekamskneftekhim" production of unblended benzine by combination of 7 components differ in density (technology of turbulent compounding) was or-

152

ganized. Non-optimal conditions of mixing didn't allow continuous production of benzine of required quality. Application of multizone tubular turbulent apparatus of divergent-convergent design (Fig. 6.10) solved the problem.

6

7

3

4

Figure 6.10. Tubular turbulent apparatus for preparation of unblended benzine (1-8 - see description text). Device has body 1 made as thick-walled cylinder in which static mixer elements in the form of combined cones are placed - convergents 3 and divergents 4, they divide the device on sections. Consecutive introduction of components and mixture mixing regime are reached at the expense of apparatus constructive particularities and allow effective compounding of fractions with various boiling temperature and with large interval of volume throughput that is of great importance when compositions of motor petrol are prepared in production quantities. Mixing of high-boiling components of fuel composition in the first section allows receipt of stable components mixture and then in subsequent sections to mix it with the rest components flows, for example with virgin benzin, benzine of catalytic cracking or refining, and by this stable in structure unblended benzine is continuously formed. This process is applied at the plant "Ethylene" of OAO "Nizhnekamskneftekhim" from 1999. 6.3.2. Alkyl-sulfates production Interaction of olefine hydrocarbons with sulfuric acid is widely used in industrial production, in particular under receipt of synthetic cleaning agents, methylethylketone, benzine fractions of hydrocarbons [I], higher fatty alcohols [276], etc.: RCH= CH2 + H2S04 + R-CH-O-SO3H + Q

I

CH3 Anionic surfactant species on the base of alkylsulfates are widely used as a base for synthetic cleaning agents, emulsifiers, stabilizers of foams, etc.. Alkylsulfates ROS03Me (Me - Na, K and so on) at conditions of industrial production, in particular under synthesis of cleaning agents are produced under interaction of a-olefines fractions of unsatureated hydrocarbons C&,S (olefins content not lower than 97 mass %) with 98% sulfuric acid at the ratio a-olefine / HzS04 = 1/1,2 + 1,4 mole with further neutralization of unreacted HzS04 and saponification of received alkyl-sulfuric acid:

R-CH-O-S03H + NaOH

I

CH3

-

153

R-CH-O-S03Na + H 2 0

I

CH3

Traditional technological scheme of alkylsulfates receipt includes following stages (Fig. 6.1 1a) [277]: sulfation of a-olefines by concentrated sulfuric acid (position la, lb); neutralization of unreacted HzSO4 and alkylsulfuric acid (position 2a-2d); two-stage cleaning by benzine of received alkylsulfates from unsaponifiable hydrocarbons products (position 4a, 4b, 5). Selection of reactor for sulfation of a-olefines caused significant complications due to specificity of chemical process proceeding. Reaction proceeds fast enough (characteristic reaction time according with [276] is lower than 30 sec) and is high-exothermal. In particular, heat of sulfation is equal to q 0 502,1.103 kJ per 1 m3 of a-olefines [277]. As a consequence temperature rise in reaction zone is about 19Ok5OC. We should also note that due to significant difference in values of viscosity (p) and density (p) of initial reagents (pH ~ S pOo l~e f i n e 7531; z 2,63) the

polefine fine

flows are mixed badly and by this especially at the initial period reaction proceeds on phases contact surface that leads to deceleration of chemical reaction proceeding in spire of the fact that this reaction may be ascribed to fast chemical processes. In this case increase of turbulent level of mixing in reaction zone leads to the rise of chemical process proceeding rate (Fig. 6.12). As a consequence under technological design of alkylsulfuric acid receipt stage in production of alkylsulfates it is necessary to form in reaction zone intensive turbulent mixing of initial reagents and to provide removal of resulted in the course of reaction heat. Moreover, presence of side processes (formation of dialkylsulfates, resinification of a-olefines, hydrolysis, and also alcoholysis of dialkylsulfates, etc.) determines the necessity of reaction mixture residence times Z~ reduction and definite limitation of upper temperature level in reaction zone (reactor).

154

f

Ib

l a

w-

benzine

+

benzine

I

I

to the slag-collecto! alkylsulfates to evaporation u-olefine

benzine, to regeneration

benzhl 4

-no,

L

y

,

to the slag-collector

U

benzinet

alkylsulfates , to evaporation b)

Figure 6.11. Technological scheme for receipt of alkylsulfates: la, l b - tank sulfators with impeller mixers (Vr = 0,157-0,44 m3); 4a, 4b - extraction columns of the first stage (vr = 3,4 m3); 5 - extraction column of the second stage (v, = 4,7 m3) (scheme a), 1 - tubular turbulent sulfator (Vr = 0,0086 m3); 4 - tubular turbulent extractor of the first stage (v, = 0,003 m3); 5 - tubular turbulent extractor of the second stage (v, = 0,005 m3); 6 - sediment box (scheme b), 2a-2d - tank reactors-neutralizator with propeller mixers (v, = 3 m3);. 3 - seoarator.

155

Yield, YO 100 80

60 40

20 0

Figure 6.12. Conversion curves of sulfation of a-olefines of fraction C&I6 by concentrated (95 mass YO)sulfuric acid in stirred tank reactors. Ratio a-olefines / H z S O ~ = 3,9 volume. 298 K. Mixing rate, revs: 100 (l), 500 (2), 1000 (3). This required intensive turbulent mixing of liquid flows in apparatus, large specific surface for improving of external heat removal, jet regime under reaction with easy controlled reagents residence time T ~ that , are possible under tha use of apparatus of novel generation - compact tubular turbulent reactors for effective carrying out of reaction of a-olefines sulfation by sulfuric acid at conditions of industrial production. One of the real ways of essential decreasing of adiabatic temperature rise in reaction zone ATad under a-olefines sulfation by sulhric acid is dilution of initial reaction mixture by inert solvent (saturated hydrocarbons, for example decane) with high heat capacity (Table 6.3) and there is also significant reduction of cooling zone length Lcoo~ (Fig. 6.13). Application of low-boiling solvent for dilution of initial reaction mixture allows effective use also of internal heat removal at the expense of its boiling. In this case gasification in reaction mixture volume allows additionally turbulization of flow and high linear rates of movement prevent from gas "locks" formation. The process becomes simple, compact, effective and easy control. Table 6.3. The influence of dilution of a-olefines by decane on heat regime of sulfation in tubular turbulent apparatus without internal heat removal (To = 253 K, Tch = 253 K, R = 0,014 m, V = 0,35 m/sec, apparatus

156

As a result of proceeding of a-olefines sulfation reaction by concentrated sulfuric acid in reaction volume heat Aq,d is evaluated: &ad = qwolefine (6.1) Here Wolefine - volume throughput of hydrocarbon fraction of a-olefines. The heat qad is spent on adiabatic heating of reaction m a s up to the temperature determined by solvent boiling temperature Tboil: qad = qWZP(Tboi1 - TO) (6.2) and using equations (6.1) and (6.2) one may obtain amount of low-boiling solvent ws required for removal of the residuary heat amount at the expense of boiling:

w = "ad S

-'ad 9,"

where qev - solvent evaporation heat.

170

50

1

! 0

I

1

I

2 decane/a-olefine. volume

Figure 6.13. Dependence of the length of cooling zone Leoolon dilution of initial reaction mixture under the use of "zone" model. Introduction of equal amounts of H2S04 in 1 (l), 2 (2), 3 (3) and 4 (4) reaction zones at the absence of internal heat removal (To = 303 K, T, = 278 K, T,h = 253 K, R = 0,014 m, wolefine = 0,62 m3/hour, w = 0,16 m3/hour). For example, under dilution of initial reaction mixture by petroleum ether with Tboil = 343 K in ratio a-olifine / solvent = 1 / 1,24 volume at the expense of adiabatic heating of reacting mixture from To = 293 K up to Tboil = 343 K (AT = 5OoC) in tubular turbulent apparatus at wolefine = 0,62 m3/hour, WHtSo4= 0,16 m3/hour and R = 0,014 m 1,5.105 kJ of heat are removed, i.e. about 48% of heat. At the expense of solvent boiling (petroleum ether) at atmosphere pressure in apparatus 1,6.105 kJ are removed, i.e. residuary 52% of heat. If one will use dilution in ratio aolefine / petroleum ether = 1 / 1 volume, so at the expense of adiabatic temperature rise and solvent boiling about 86% of heat will be removed and temperature in sulfator will rise at the expense of auto-thermal heating in 1517°C higher than boiling of solvent that is permissible. In the last case after removal of the main petroleum ether part out of gas phase in real production the scheme of alkyl-sulfuric acid purification from solvent's remains is simplified due the absence of the last one.

157

The whole technological scheme is also simplified on the base of compact tubular turbulent apparatus (Fig. 6.1 lb) in comparison with the scheme existing in production now (Fig. 6.1 la). There is also alternative method of effective decrease of adiabatic temperature rise at points of reagents introduction - "zone" model of process carrying out (see 4.2.1, Fig. 6.ld) at the expense of distribution of given amount of sulfuric acid along the length of tubular turbulent apparatus of cylindrical design. In particular increase of number of zones of portion introduction of lower amount of H2S04at introduction points allows sharp reduction of ATad and preventing of sharp temperature rise in reaction zone higher than technologically acceptable value (Fig. 6.14). As a consequence, under increase of number of reagents introduction zones in tubular apparatus "deceleration" of fast reaction is observed expressed in the fact that tempetature profile in reactor approaching to chemical process proceeding at &h = T~~ in ideal mixing regime. In this case even at the absence of external heat removal reaction mixture temperature over the whole apparatus length is equal, i.e. process proceeds in quasi-isothermal regime. However, under the use of "zone" model of process carrying out the technologic scheme becomes complicated, in particular at the expense of necessity of use of several cooling devices between reaction zones and increase of summary length of cooling zone Leooland consequently reactor's dimensions (Fig. 6.14).

500

400

'*.

300

5..

\,*I..

A 'm 2

200

! 0

1

50

'.

'm 3 1

looL , m 150

Figure 6.14. Temperature profile T along tubular turbulent sulfator length L under the use of "zone" model. Introduction of equal amounts of into 1 (l), 2 ( 2 ) , and 4 (3) reaction zones (condition as in Figure 6.13). For simplification of process' technological line under the use if zone model, in particular under sulfation of a-olefines by sulfuric acid at the cooling stage it is advisable to use shell-and-tube turbulent apparatus of jet type including pipe bundle of small diameter washing by coolant (Fig. 6 . 1 ~ ) Fragmentation . of flow into N tubes of smaller radius r and working in quasi-plug flow regime at maintaining of general section of flow S at in single tubular apparatus under the use of both traditional and zone models allows reduction of cooling zone length Lcoolin NOs6 times (Fig. 4.14) and making the process more compact. Combination of portion introduction of sulfuric acid and application of multi-tubing cooling inserts allow additional decreasing of temperature adiabatic rise ATad and significant reduction of cooling zone length under sulfation of a-olefines (Fig. 6.15).

158

350

300

250 0

10

20

L, m

30

Figure 6.15. Temperature profile along tubular turbulent sulfator length under two-stage introduction of equal amounts of HzS04. Reaction zone (1) - tubular turbulent apparatus of cylindrical design (R= 0,014 m); cooling zone ( 2 ) - shell-and-tube turbulent apparatus (N = 31, r = 0,0025 m). To = 253 K, T, = 278 K, Tch = 253 K, Wolefine = 0,62 m3/hour, W, so = 0,16 m3/hour). 2

4

Thus, reaction of a-olefines sulfation by concentrated sulfuric acid is related to fast chemical reactions that are necessary to carry out in compact energy- and resource-saving tubular turbulent apparatus. For guarantee prevention of sharp temperature rise in reaction zone under sulfation of aolefines application of tubular turbulent apparatus of cylindrical and shell-and tubes designs is advisable. By this one should dilute initial reagents especially by low-boiling solvent and realize "zone" model of process carrying out (portion reagents introduction along reactor length). Novel energy- and resource-saving technological scheme of alkylsulfates receipt at conditions of industrial production with the use of compact tubular turbulent apparatus at the stage of a-olefines sulfation by sulfuric acid and washing of resulted alkylsulfates was proposed. In 200 1 recommendations on application of tubular turbulent apparatus of divergentconvergent design at two stages of technological scheme of production of cleaning agent "Progress" (Fig. 6.1 1b) were given for PhSUE "Novocherkassky plant of synthetic products" (Novocherkassk city, Rostovsk region): at stages of sulfation of cracking-product 1 (vp 0,0086 m3) and extraction of alkylsulfates from unsaponifiable hydrocarbons (the first stage 4 - vp 0,003 m3, the second stage 5 - vp 0,005 m3). At present designing and preparation of novel experimental-industrial technological scheme of alkylsulfates production with output equal to 300 tins per year with the use of tubular turbulent apparatus are completed. They expect that at the expense of more effective reagents mixing by novel technology in comparison with basic on the base of stirred tank reactors the yield of desired product with realization of soft reaction proceeding regime will increase.

-

-

-

6.3.3. Neutralization of acid and alkali mediums

Stage of neutralization of acid mediums in liquid-phase flows is integral part of the majority of chemical processes of synthetic products production in industry including polymers synthesis. As the damping of industrial wastewaters to sewerage system is permitted at pH 2 7, then stage of neutralization of acid mediums is the key stage under purification of industrial effluents and also under receipt of many synthetic products (neutralization of aminochlorohydrates under ehtylenediamine

159

production, ammonation of phosphoric acid under production of ammophos, neutralization and saponification of alkylsulfuric acid under alkylsulfates production, washing out of inverse solvent in polymers production, etc.). Under neutralization stage carrying out at conditions of industrial production it is necessary to take into account specific particularities of this process. Neutralization reaction is very fast (length of reaction zone is very small L c h = 2.10-'' m (Table 4.1). This caused its proceeding in classic stirred tank reactors in diffusion regime even at low reagents concentrations, by this neutralization reaction is highly exothermal process (heat effect of interaction of strong acid with strong base is q = 57,2 kJ/mole). Process proceeds in aggressive mediums that at high temperature require anti-corrosion apparatus design. Under process optimization effective mixing of initial reagents solvents for the time ~ , , , i commensurable ~ with the time of chemical reaction Tch, i.e. at rmix = Tch and also application of apparatus without moving mechanical devices with small specific area of reaction mixture with reactor's walls contact and decreasing of sharp temperature rise in neutralization agent introduction zone are required. Experience of application of apparatus of novel generation - compact tubular reactors functioning in regime of quasi-plug flow mode in turbulent flows without application of mixers for fast and many of mass-exchange physical processes in combination with results of laboratory research and mathematical modelling of liquid flows mixing processes (see chapter 3) allow us to suggest their effective functioning under neutralization of acid effluents that can not be reached under the use of any other known apparatus of applied chemistry. In this case it is important to reveal possibilities of formation of plan reaction front in reaction zone corresponding to quasi-plug flow mode in turbulent flows and providing quasi-isothermal regime in reaction zone (see Chapter 4). Obtained experimental results allow derivation of dependence of conditions of quasi-plug flow mode formation in tubular turbulent apparatus at VI, V2 = 0,1+0,8 m/sec and dl / dd = 0,44: V l N 2 = 0,28+V2+ 0,96 (6.4) Since the flows differ in density and viscosity are introduced for neutralization application of dependences (4.2) and (4.3) allows carrying out of the process in conditions of industrial production in quasi-plug flow mode in turbulent flows. Using (1.4) one can estimate the fact that under neutralization of 30% water solution of HCI by NaOH water solution temperature rise AT,d is equal to 50"C, at the same time under neutralization of 30% acid solution by solid alkali AT,d will be significantly higher and about 130°C (Table 6.4). Table 6.4 Temperature rise in the course of neutralization of water solution of hydrochloric acid by sodium hydroxide.

It is obvious from presented results that process of neutralization of acid mediums (to the extent of 30% acid water solution) may be carried out by water solution of alkali (up to 20 mass %) without additional external heat removal at initial temperature of acid effluent not higher than 30°C. in this case temperature rise ATad in reaction zone will not exceed 30-50°C (Table 6.4). Under the use of reagents concentrated solutions for neutralization of acid mediums sharp temperature rise is observed. In this case the novel type of apparatus of applied chemistry - com-

160

pact tubular turbulent reactors of cylindrical or divergent-convergent designs determine the possibility of effective control of temperature field in reaction zone by several ways (see 4.2): change of apparatus radius and reagents flow rate, application of "zone" model for fast chemical process carrying out and the use of shell-and-tube apparatus with pipe bundle of N tubes of small radiuses. It should be noted that in the necessity of effective mixing of reagents and preventing of possible acid slip caused by the fact that real flow of industrial effluent significantly exceeds the consumption of neutralization agent it is advisable to carry out the reaction in tubular turbulent apparatus of divergent-convergent design. In this case one should apply reactor with diameter in the wide part (divergent) from 0,05 up to 0,4 m and length up to 2 m (apparatus volume is about 0,016-0,12 m3) that corresponds the productivity of one compact reactor about 7 m3/hour - 460 m3/hour and higher. The holes in socket for alkali (acid) solution introduction into reactor (placed coaxialy to the flow) should provide reagents introduction into reaction zone with the rate VI > 1,l m/sec to provide the ratio VI / V2 2 1,2. Tubular turbulent apparatus of divergent-convergent design at present are used in industrial production under neutralization at stages of isoprene hydrochlorides destruction in production of cis-l,4-isoprene rubber of SKI-3 and SKI-5 brands ( A 0 "Kauchuk", Sterlitamak city), Ziegler-Natta catalysts decomposition (OAO "Nizhekamskneftekhim"), etc.. Thus, reaction of neutralization of acid and alkali mediums relates to fast chemical processes which are advisable to carry out in reactors of novel type - compact energy- and resource-saving tubular turbulent apparatus of cylindrical or divergent-convergent designs that provides jet regime of reagents flow, quasi-isothermal conditions in reaction zone, total conversion of reagents, etc.. Neutralization of industrial effluent with mass content of acid up to 30 mass % can be realized in tubular turbulent apparatus by both 10% and 20% water alkali solution without additional heat removal. Results obtained in this work may be also used under realization of neutralization of alkali effluent by acids.

161

CONCLUSION Under realization of fast processes (chemical or heat-mass-transfer) in conditions of industrial production the main problem is the necessity of intensive turbulent mixing in reaction zone with the aim of equalization of distributions by reagents concentrations and temperature, Le. creation of uniform synthesis conditions. The optimal solution of this significant scientific and applied problem is to carry out fast processes in turbulent flows - in compact tubular turbulent apparatus. Works in this area are at the turn of chemical physics and applied chemistry that requires comprehensive approach to studying of physical-chemical basis of fast processes proceeding. From the one hand, search of effective methods of formation of high level of turbulent reagents mixing and also removal of significant amount of locally evolving heat are required and allow indirect influence on resulted synthetic products quality. From the other hand, investigation of specificity of concrete processes proceeding and also effect of increase of turbulent mixing level in reaction zone on the character of process proceeding and quality of resulted products are necessary. This approach allows working out of novel solutions in the field of technological processes apparatus design under fast chemical reactions realization. Apparatus of divergent-convergent design is the optimal in the view of diffusion limitations relieving under fast processes proceeding. This is determined by the fact that during increasing of divergent opening angle y from 5’ up to 30°, i.e. during transition from cylindrical to divergentconvergent design turbulent diffusion coefficient in reaction zone rises practically in 3 times. High level of turbulent mixing is created only at the expense of geometry of the canal itself and remains constant as moving off reagents introduction point. As a consequence, divergent-convergent design in comparison with cylindrical allows significant broadening of tubular turbulent apparatus working regimes to the region of low flows rates, i.e. larger reagents residence times and slower reactions. The ratios dd / d, = 1,6 and L, / d, = 1,7 are optimal in relation to minimum time of reaction mixture micro-mixing in reaction zone of divergent-convergent design. In cylindrical canal with 4 10,03 m the regime is formed close to quasi-plug flow mode that is connected with low rate of longitudinal mixing and narrow distribution by reagents residence times in reaction zone. Because of high rate of longitudinal mixing and hence the wide distribution by reagents residence times the regime of ideal mixing is formed in apparatus of divergent-convergent design that allows realization of isothermal conditions in plug-flow reactor. Under radial and coaxial reagents introduction in one and the same reaction volume in dependence on the ratio of flows introduction rates the formation of regimes with various approximation degree both to idealized mixing and displacement models is possible that is impossible in existing reactors of applied chemistry. This is connected with the possibility of formation of characteristic macroscopic structures in reaction zone (plan front, torch, drift and so on) in dependence on reagents introduction rates ratio. Reactor of divergent-convergent design allows significant broadening of the possibilities of continuous tubular apparatus in relation to viscosity of reaction mixture. Particularly, the criterion Re 2 Recr= 950+50 was proposed and the possibility of formation of self-similar regime of liquid flows was revealed when viscosity affect on turbulent mixing characteristics was levelled. Low value of Recr in reaction zone of divergent-convergent design is determined by the fact that canal geometry at the expense of generation of hydrodynamic disturbances provides formation of large-scale turbulent pulsations already at low rates of reaction mixture movement (for cylindrical canal Recr- lo’). Possibility of formation of self-similar region allowed receipt of analytical expressions for calculation of average values of turbulent diffusion coefficient Dt, turbulence kinetic energy density K and the rate of its dissipation E, and also of characteristic times of mixing at various levels. Fundamental bases and regularities of multi-phase and laminated systems flows in apparatus of divergent-convergent design were worked out for the first time. Analytical dependences for calculation of turbulence kinetic energy density dissipation rate E of disperse medium and maximum -0,4 value of dispersion phase particles diameter d,, = 0,099(a / pl)06 ’ E under reaction mixture flow in turbulent regime were obtained. Geometry of reaction zone significantly influences on specific interface of reacting phases. In particular, increase of the number of divergent-convergent sections

162

from 1 up to 4 leads to the decrease of diameters of dispersion insertions that allows application of apparatus with the length 8-10 calibers. Increase of the ratio dd I d, (transition from cylindrical to divergent-convergent canal) determines the reduction of disperse systems movement rate required for formation of flow with uniform distribution of disperse phase. Uniform flow with homogeneous distribution of disperse phase particles in reaction volume of divergent-convergent design is observed in the interval of disperse system linear rate 0,l < V < 16 d s e c . Bottom limit is limited by reaction mixture lamination due to difference in reagents densities, and limit superior - due to separation effect in peripheral apparatus part at the expense of centrifugal forces. Analytic formula for estimation of components separation degree of multi-phase reaction mixture at high rates of flows movement Act2 I aza,= 0,06Ar''44I Re2'' was proposed. This allows formation of homogeneous reaction mixture under fast chemical reactions proceeding in synthesis of polymers at phases interface. In general case, analysis of physical picture of turbulent movement of homogeneous and heterogeneous reaction mixtures in tubular canals of various geometry determined the possibility of quantitative and scientifically-grounded approach to selection of optimal geometry of reaction zone under proceeding of any fast chemical processes with the aim of decreasing of diffusion limitations and creation of homogeneous conditions under polymers synthesis. Fundamental and applied aspects of fast polymerization processes proceeding allowing the solution of large problem of reaction zone temperature field control at the expense of external heat removal that was impossible earlier were developed. Quantitative regularities of conditions of quasi-plug flow mode formation determining effectiveness of external heat removal of fast chemical processes were obtained in dependence on kinetic parameters of proceeding reaction and also physical characteristics of reacting flows. Formulas for calculation of temperature profiles in reaction zone under quasi-isothermal regime formation at conditions of external heat removal were received and their adequacy was experimentally confirmed. On this base technologically acceptable methods of heat regime control under fast exothermal reactions proceeding particularly by the change of reaction zone radius and reagents movement rate, of realization of "zone" model of process carrying out, application of shell-and-tube design with tubes bundle of small radius and also change of reaction mixture flow hydrodynamic regime were revealed. The use of reaction zone of divergent-convergent design allows increasing in 1,4-1,7 times of coefficient of heat-transfer through the wall and in 1,3-1,4 times of apparatus specific productivity. This occurs at the expense of the fact that reaction zone of divergent-convergent design in relation to hydrodynamic structure of reaction mixture movement is characterized by high approximation degree to ideal mixing regime with intensive longitudinal heat transfer that significantly intensifies convective heat-exchange. Fundamental regularities of increase of turbulent mixing in reaction zone and heat regime control received in this work allow significant broadening of tubular turbulent apparatus application fields under polymers synthesis including intensification of heat- and mass-transfer processes. Possibility of realization in turbulent regime not only fast polymerization processes, but also polymer-analogous reactions under polymers modifications was revealed. Possibility of formation of self-similar regime in reaction zone of divergent-convergent design solves the problem of work with high-viscous polymers solutions (butyl rubber and chloro-butyl rubber) and allows decreasing of scattering by chlorine content practically in 3 times at the expense of disperse system formation with constantly renewable surface. Conception of use of compact, high-productive, energy- and resource-saving tubular turbulent apparatus was turned to be acceptable not only for novel class of chemical processes - fast liquid-phase reactions always proceeding in diffusion regime. Differences in kinetic parameters under proceeding of one and the same chemical reaction in the presence of catalysts differing in activity determine the necessity of various approach to the selection of type and construction of the main reactor, that was demonstrated by the example of piperylene cationic polymerization. The most commonly used catalysts of piperylene polymerization according with their kinetic activity are settled in the raw TiCL < TiC14Al(i-C4H9)3 < A I C ~ H ~ C ~ ~ ~ O < ( C ~AlC2HjC12 H S ) ~ < AlC13~O(C&)2 (kp= 2+420 llmolemin)

163

without significant differences in oligomer's molecular characteristics (Mn = 880+1080; M, =1350+1890; M, / Mn = 1,5+1,8).Kinetic analysis showed that under transition from Tic14 to AlC13.O(C6&)2 necessary reaction zone size was decreased and realization of quasi-isothermal conditions required reduction of apparatus dimensions. Interesting and practically important idea of separation of processes of olefins and dienes (co)polymerization in the presence of complex Ziegler-Natta catalysts into fast (formation of active sites and reaction mixture) and slow ((co)polymerization itself) stages was proposed. In practical aspect it determines advisability of use of tubular turbulent pre-reactor before stirred tank reactor with large reagents residence time in reaction zone for realization of the main process ((co)polymerization). The novel and perspective field of tubular turbulent apparatus application is the carrying out of fast chemical processes with formation of solid phase. This provides wide possibilities for modification of metal-complex Ziegler-Natta catalytic systems at the expense of hydrodynamic influence on disperse structure of micro-heterogeneous catalysts particles, receipt of high-effective anti-agglomeratesfor synthetic rubbers, etc.. Thus, decrease of diffusion limitations under polymers synthesis at the expense of turbulent mixing intensification in reaction zone allows calculation of kinetic parameters of polymerization and polymer-analogous processes, optimization of molecular characteristics of resulted polymer products and to influence on macro-kinetic particularities of the process in the whole. Change of geometry (design) of mixing zone of turbulent flows limited by impenetrable wall, physical parameters and reagents dynamics allow easy control of processes of mass and energy transfer in accordance with specificity of proceeding process. Physico-chemical basis of fast processes realization in turbulent regime received in this work serves as a base for creation of a number of energy- and resource-saving high-productive technologies on the base of novel type of reactors of applied chemistry - tubular apparatus of jet type.

164

REFERENCES

[l] AI. AI. Berlin, K. S. Minsker, K. M. Dyumaev New UniJied Energy- and Material-Saving, High-Performance, Ecological-Friendly Technologies based on Tubular Turbulent Reactors, Moscow (1996) (in Russian). [2] AI. AI. Berlin, K. S. Minsker Technology of the XXI century. Nauka proizvodstvu, 53, No. 3, 7-12 (2002). [3] A. G. Kasatkin Basic Processes and Apparatuses in Chemical Technology, Khimiya, Moscow (1971) (in Russian). [4] K. Denbig Theory of chemical reactors, Nauka, Moscow (1968) (in Russian). [5] Kh. Kramers, K. Vesterterp Chemical reactors, Nauka, Moscow (1967) (in Russian).

[6] Yu. I. Dytnerskii Processes and Apparatuses in Chemical Technology, Part 1, Khimiya, Moscow (1995) (in Russian). [7] D. A. Baizenberger, D. Kh. Sebastian Principles of Polymerization Engineering, Khimiya, Moscow (1988) (in Russian). [8] AI. AI. Berlin, S. A. Vol'fson Kinetic Methods in Polymers Synthesis, Khimiya, Moscow (1973) (in Russian). [9] N. H. Kolphapure, R. 0. Fox Chemical Engineering Science, 54,3233-3242 (1999). [lo] AI. AI. Berlin, S. A. Vol'fson, N. S. Enikolopyan Kinetics of Polymerization Processes, Khimiya. Moscow (1978) (in Russian). [ l l ] AI. AI. Berlin, S. A. Vol'fson Vysokomol. Soedin., 36, No. 4, 616-628 (1994). [ 121 Z. M. Tzareva, L. L. Tovazhnyanskii, E.I. Orlova Basis of Chemical Reactors Theory, Kharkov

(1997) (in Russian). [13] V. V. Kafarov Cybernetics Methods in Chemistry and Chemical Technology, Khimiya, Moscow (1985) (in Russian). [14] J. Villermaux Chemical Engineering Science, 51,No. 10, 1931-1946 (1996). [15] A. B. Golovanchikov, G. V. Ryabchuk, N. A. Dul'kina Izvestiya VUZov. Khimiya i Khimicheskaya Tekhnologiya,44,No. 4, 87-89 (2001). [16] V. V. Kafarov, M. B. Glebov Mathematical Modeling of Basic Chemical Productions Processes, Vysshaya shkola, Moscow (1991) (in Russian). [17] A. P. Torres, F. A. R. Oliveira, S. P. Fortuna Journal of Food Engineering, 35, No. 2, 147-163 (1998). [I81 V. M. Ramm, Gases Absorption, Khimiya, Moscow (1976) (in Russian).

165

[19] A. M. Zhurba, S. V. Kurilov Theoretical basis of applied chemistry, 25, No. 3,440-444 (1991). [20] A. D. Martin Chemical Engineering Science, 55, 5907-5917 (2000). [21] M. M. Gavrilenko, I. E. Maleev, V. I. Nikitin Khimiya i Khimicheskaya Tekhnologiya, 41, No. 3, 101-104 (1998). [22] A. M. Zhurba, S. V. Kurilov, L. L. Gerasimova, I. B. Vil'chinskaya Theoretical basis of applied chemistry, 29, No. 1,22-30 (1995). [23] 0. Levenshpil' Engeenering Design of Chemical Processes, Khimiya, Moscow (1969) (in Russian). [24] P. Viitanen Applaid Radiation Isotopy, 48, No. 7, 893-898 (1997). [25] K. S. Minsker, Al. AI. Berlin, V. P. Zakharov Vysokomol. Soedin., 44, No. 9, 1606-1627 (2002). [26] Yu. A. Sangalov, K. S. Minsker Polymers and Copolymers of Isobutylene: Fundamental Problems and Applied Aspects, Gilem, Ufa (2001) (in Russian). [27] K. S. Minsker, Yu. A. Sangalov Isobutylene and its polymers, Khimiya, Moscow (1986) (in Russian). [28] Al. Al. Berlin, K. S . Minsker, Yu. A. Prochukhan, M. M. Karpasas, N. S. Enikolopyan Vysokomol. Soedin., 28, No. 6,461-465 (1986). [29] V. Z. Kompanietz, A. A. Konoplev, Al. AI. Berlin, Yu. A. Prochukhan, K. S. Minsker, M. M. Karpasas, N. S. Enikolopyan Dokl. Akad. Nauk, 297, No. 5, 1129-1132 (1987). [30] A. Ya. Fedorov, G. E. Litvak, L. P. Kholpanov, V. A. Malyusov Dokl. Akad. Nauk, 319, No. 2, 422-426 (199 1). [31] A. Ya. Fedorov, B. L. Rytov, Al. Al. Berlin, G. G. Aleksanyan Dokl. Akad. Nauk, 342, No. 4, 494-498 (1995). 1321 Al. AI. Berlin, K. S. Minsker, Yu. A. Prochukhan, N. S . Enikolopyan Vysokomol. Soedin., 31A, NO. 9, 1779-1798 (1989). [331 AI. AI. Berlin, K. S. Minsker, Yu. A. Prochukhan, N. S. Enicolopyan Polym. Plast. Technolog. Eng., 2-3,253-297 (1991). [MI AI. Al. Berlin, K. S. Minsker, Yu. A. Sangalov, D. A. Novikov, T. N. Poznyak, Yu. A. Prochukhan, A. I. Kirillov, A. G. Svinukhov Vysokomol. Soedin., 21B, No. 6,468-471 (1979). [35] AI. Al. Berlin, K. S. Minsker, Yu. A. Prochukhan, M. M. Karpasas, N. S. Enikolopyan Dokl. Akad. Nauk, 287, No. 1, 145-148 (1986). 1361 Ya. B. Zel'dovich Chemical Physics and Hydrodynamics, Nauka, Moscow (1984) (in Russian). [37] V. Z . Kompaniets, A. A. Ovsyannikov, A. S . Polak Chemical Reactions in Turbulent Flows of Gases and Plasma, Nauka, Moscow (1979) (in Russian).

166

[38] AI. AI. Berlin, K. S. Minsker, Yu. A. Sangalov, V. G. Oshmyan, A. G. Svinukhov, A. P. Kirillov, N. S. Enikolopyan Vysokomol. Soedin., 22A, No. 3, 566-574 (1980). [39] K. S. Minsker, Al. AI. Berlin, A. G. Svinukhov, Yu. A. Prochukhan, N. S. Enikolopyan Dokl. Akad. Nauk, 286, No. 5 , 1171-1173 (1986). [40] Al. AI. Berlin, K. S. Minsker Dokl. Akad. Nauk, 355, No. 3, 346-348 (1997). [41] AI. Al. Berlin, K. S. Minsker, K. M. Dyumaev, S. V. Kolesov, S. P. Gantseva Khimicheskaya Promyshlennost', No. 3,54-57 (1997). [42] S. K. Minsker, T. V. Golubeva, A. A. Konoplev, V. Z. Kompaniets, Al. AI. Berlin, K. S. Minsker, N. S. Enikolopyan Dokl. Akad. Nauk SSSR, 314, No. 6, 1450-1454 (1990). [43] S. K. Minsker, A. A. Konoplev, K. S. Minsker, Yu. A. Prochukhan, V. Z. Kompaniets, Al. AI. Berlin Theoretical basis of applied chemistry, 26, No. 5,686-691 (1992). [44] K. S. Minsker, K. M. Dyumaev, AI. Al. Berlin, N. P. Petrova, S. K. Minsker, A. Ya. Fedorov Bashkir. Khim. Zh., 2, No. 3-4,41-45 (1995). [45] AI. Al. Berlin, K. S. Minsker, K. M. Dyumaev, S. V. Kolesov, S. P. Gantseva Khimicheskaya Promyshlennost', No. 5,27-30 (1997). [46] D. A. Frank-Kamenetskii Diffusion and Heat-Transfer in Chemical Kinetics, Nauka, Moscow (1987) (in Russian). [47] AI. Al. Berlin Sorosovskii Obrazovatelnyi Zh., No. 3,48-54 (1998). [48] A. Ya. Fedorov, G. E. Litvak Vysokomol. Soedin., 33A, No. 10,2626-2634 (1991). [49] AI. AI. Berlin, G. E. Zaikov, K. S. Minsker Polymer News, 22, No. 8, 291-292 (1997). [50] AI. AI. Berlin, K. M. Dyumaev, K. S. Minsker, F. R. Khalafov, S. V. Kolesov Khimicheskaya Promyshlennost', No. 9,550-556 (1995). [51] K. S. Minsker, AI. Al. Berlin, Yu. A. Prochukhan, N. S. Enikolopyan Vysokomol. Soedin., 28B, NO. 6,466-469 (1986). [52] K. S. Minsker, AI. AI. Berlin, Yu. A. Prochukhan, N. S. Enikolopyan Vysokomol. Soedin., 28B, NO. 6,466-469 (1986). [53] G. Khemmis Methods offast reactions studying, Mir, Moscow (1977) (in Russian). [54] Yu. A. Prochukhan, AI. Al. Berlin, K. S. Minsker, N. S.Enikolopyan Dokl. Akad. Nauk SSSR, 287, NO.3, 682-685 (1986). [55] AI. AI. Berlin, K. S. Minsker, R. Ya. Deberdeev Dokl. Akad. Nauk, 375, No. 2, 218-221 (2000).

167

[56] AI. AI. Berlin, V. Z. Kompaniets, A. A. Konoplev, K. S. Minsker, S. K. Minsker, Yu. A. Prochukhan, E. A. Ryabenko, N. S. Enikolopyan Dokl. Akad. Nauk SSSR, 305, No. 5, 11431146 (1989). [57] X. Wang, Z. Feng, L. J. Fomey Computers and Structures, 70,447-465 (1999). [58] Yu. A. Prochukhan, K. S. Minsker, AI. AI. Berlin, M. M. Karpasas, V. Z . Kompaniets, A. A. Konoplev, N. S. Enikolopyan Dokl. Akad. Nauk SSSR, 298, No. 6, 1428-1430 (1988). [59] Yu. A. Prochukhan, K. S. Minsker, M. M. Karpasas, AI. AI. Berlin, R. Kh. Bakhitova, N. S. Enikolopyan Vysokomol. Soedin., 30A, No. 6, 1259-1262 (1988). [60] AI. AI. Berlin, Yu. A. Prochukhan, K. S. Minsker, E. A. Tumanyan, G. G. Aleksanyan, N. S. Enikolopyan Vysokomol. Soedin., 30A, No. 11,2436-2440 (1988). [61] K. S. Minsker, AI. AI. Berlin, Yu. A Prochukhan., E. A. Tumanyan, M. M. Karpasas, N. S. Enikolopyan Dokl. Akad. Nauk SSSR, 291, No. 1, 114-116 (1986). [62] AI. AI. Berlin, Yu. A. Prochukhan, K. S. Minsker, N. S. Enikolopyan Vysokomol. Soedin., 33A, NO. 2,243-269 (1991). [63] Yu. A. Prochukhan, K. S. Minsker, AI. AI. Berlin, E. A. Turnanyan, N. S. Enikolopyan Dokl. Akad. Nauk SSSR, 291, No. 6,1425-1427 (1986). [64] AI. AI. Berlin, Yu. A. Prochukhan, K. S. Minsker, G. G. Aleksanyan, S. V. Grobov, N. S. Enikolopyan Vysokomol. Soedin., 31A, No. 3, 612-616 (1989). [65] AI. AI. Berlin, Yu. A. Prochukhan, K. S. Minsker, G. G. Aleksanyan, S. V. Grobov, N. S. Enikolopyan Vysokomol. Soedin., 30A, No. 11,2441-2446 (1988). [66] K. S. Minsker, AI. AI. Berlin Fast Polymerization Processes, Gordon and Breach Publ. (1996). [67] P. D. Sarkisov Khimicheskaya Promyshlennost', No. 1,20-27 (2000). [68] K. S. Minsker, AI. AI. Berlin, Ya. M. Abdrashitov, Yu. K. Dmitriev, V. D. Shapovalov, V. P. Malinskaya, N. V. Alekseeva Khimiclzeskaya Promyshlennost', No. 6,41-44 (1999). [69] Pat 5397179 USA Al. Al. Berlin, Yu. A. Prochukhan, K. S. Minsker, A. A. Konoplyov, V. Z . Kompanietz (1995). [70] AI. AI. Berlin, K. V. Prokofev, K. S. Minsker, Yu. A. Prochukhan, S. V. Kotov, B. F. Bulakov, V. A. Yasinenko, T. I. Naumova Khimiya i teklznologiya topliv i masel, No. 7, 8-9 (1988). [71]. S. V. Kotov, AI. AI. Berlin, K. V. Prokofev, K. S. Minsker, Yu. A. Sangalov, Yu. A. Prochukhan, N. A. Adilov, G. G. Aleksanyan, V. A. Yasinenko Khimiya i tekhnologiya topliv i masel, No. 6, 10-11 (1990). [72] K. S. Minsker, Yu. A. Prochukhan, S. V. Kolesov, AI. AI. Berlin, S. V. Kotov, I. Kh. Kutluguzhina, K. V. Prokofev, V. G. Kozlov, G. G. Aleksanyan Neftepererabotka i nefrekhimiya, No. 11,40-42 (1989).

168

E731 S. V. Kotov, K. V. Prokofev, K. S. Minsker, Yu. A. Sangalov, Al. Al. Berlin Khimiya i tekhnologiya topliv i masel, No. 4, 14-15 (1990). [74] S. V. Kotov, V. E. Atmandzhev, K. S. Minsker, V. A. Yasinenko, K. V. Prokofev, Al. Al. Berlin Nefepererabotka i nefekhimiya, No. 1,38-41 (1992). [75] Patent 2059656 Russia Yu. A. Prochukhan, K. S. Minsker, R. Z. Biglova, R. Kh. Rakhimov, Yu. P. Bazhenov, S. K. Minsker, I. Sh. Nasyrov, P. I. Kutuzov, Al. Al. Berlin, V. I. Vizhnyaev (1996). [76] Patent 2059657 Russia K. S. Minsker, Yu. A. Prochukhan, R. Z. Biglova, R. Kh. Rakhimov, Yu. P. Bazhenov, S. K. Minsker, I. Sh. Nasyrov, P. I. Kutuzov, S. V. Kolesov, V. I. Vizhnyaev (1996). [77] R. G. Takhavutdinov, A. G. Mukhametzyanova, G. S. Dyakonov, K. S. Minsker, Al. Al. Berlin Vysokomol. Soedin., 44A, No. 7, 1094-1100 (2002). [78] V. M. Busygin, L. M. Kurochkin, T. G. Burganov, N. R. Gil'mutdinov, Minsker K.S., G. S. D'yakonov, R. Ya. Deberdeev, Al. Al. Berlin Khimicheskaya Promyshlennost', No. 2,3-4 (2001). [79] V. M. Busygin, G. S. D'yakonov, K. S. Minsker, Al. Al. Berlin Summa Tekhnologii, No. 4, 4849 (2000). [80] Patent 2059661 Russia K. S. Minsker, Yu. A Prochukhan., R. Z. Biglova, R. Kh. Rakhimov,

Yu. P. Bazhenov, S. K. Minsker, I. Sh. Nasyrov, M. G. Krekhova, P. I. Kutuzov, V. I. Vizhnyaev (1996).

[81] G. Ver Strate, G. Cozewith, S Ju. Macromolecules, 21, No. 12,3360-3371 (1988). [82] R. N. Gimaev, Yu. A. Prochukhan, F. Kh. Kudasheva, M. A. Tsadkin, A. D. Badikova Khimiya i tekhnologiya topliv i masel, No. 5,42-43 (1998). [83] R. N. Gimaev, M. A. Tsadkin, V. M. Rakitskii, M. M. Kalimullin, F. Kh. Kudasheva, S. V. Kolesov, R. R. Rakhmanov, A. D. Badikova, G. G. Telyashev, Yu. M. Makrimenko Bashkir. Khim. Zh., 5, No. 2,59-61 (1998). [84] K. Yu. Prochukhan, E. R. Islamov, I. V. Nefedova, R. N. Gimaev, Yu. A. Prochukhan, P. G. Navalikhin, G. G. Aleksanyan Khimiya i tekhnologiya topliv i masel, No. 2, 16-17 (1999). [85] E. R. Islamov, K. Yu. Prochukhan, I. V. Nefedova, R. N. Gimaev, P. G. Navalikhin, Yu. A. Prochukhan Khimiya i tekhnologiya topliv i masel, No. 4, 12-14 (1999). [86] Patent 9504019 WO G. G. Aleksanyan, N. B. Librovich, Yu. A. Prochukhan (1995). [87] S. V. Kolesov, M. A. Tsadkin, A. D. Badikova, R. R. Rakhmanov, F. Kh. Kudasheva, R. N. Gimaev Khimiya i tekhnologiya topliv i masel, No. 4, 15-17 (2002). [88] S . I. Dvoretskii, V. V. Karnishev Izvastiya VUZov. Khimiya i Khimicheskaya Tekhnologiya, 42, NO. 3, 101-105 (1999). [89] V. I. Bodrov, S. I. Dvoretskii, D. S. Dvoretskii Teoreticheskie Osnovy Khimicheskoi Tekhnologii, 31, No. 5,542-548 (1997).

169

[90] S. I. Dvoretskii, V. V. Karnishev, D. S. Dvoretskii Khimicheskoe i neftegasovoe mashinistroenie, No. 4,4-7 (1998). [91] A. M. Mazgarov, A. F. Vil'danov, V. M. Medem, I. A. Arkhireeva, S. A. Gorokhova, B. V. Kizhaev Khimiya i tekhnologiya topliv i masel, No. 11,21-23 (1987). [92] S. A. Borisenkova, A. M. Mazgarov, A. F. Vil'danov Ross. Khim. Zh. (Khiniicheskoe obshestvo im. D. I. Mendeleeva), 39, No. 5,87-101 (1995). [93] A. M. Mazgarov, A. F. Vil'danov, S. N. Sukhov, N. G. Bazhenova, G. B. Nizamutdinova, S. F. Shiamanna, R. M. Ormiston Khimiya i tekhnologiya topliv i masel, No. 6, 11-12 (1996). [94]. A. M. Mazgarov, A. F. Vil'danov, S. F. Shiamanna, L. W. Jossens Proceedings of the 15 th Word Petroleum Congress, 877-881 (1998). [95] A. M. Mazgarov, A. F. Vil'danov, S . F. Shiamanna Journal of the Iranian Petroleuni Institute, NO. 51, 10-15 (1999). [96] A. M. Mazgarov, A. F. Vil'danov Nafekhimiya, 39, No. 5 , 371-378 (1999). [97] Utility Model 11098 R. Ya. Deberdeev, L. S. Sakhipov, V. M. Shatilov, V. S. Trifonov, A. Sh. Ziyatdinov, A. M. Ekimova, Kh. Kh. Gil'manov, V. Z. Akhmetzyanov, K. S. Minsker, G. S. Dyakonov, N. I. Ukhov (1999). [98] Patent 2141973 Russia Z. A. Abzalin, K. S. Minsker, AI. Al. Berlin., R. Ya. Deberdeev, E. I. Nagumanova, R. G. Galiev, Kh. V. Mustafin, L. M. Kurochkin, N. R. Gil'mutdinov, I. Ryazanov, N. I. Ukhov, V. P. Pogrebtsov, T. G. Burganov, G. V. Baev, A. I. Vorob'yev, R. G. Salakhutdinov (1999). [99] K. S. Minsker, V. P. Zakharov, A.D. Ishtiryakov, 0. V. Sofronova, R. Ya. Deberdeev, G. S. Dyakonov, R. T. Shiyapov, N. R. Gil'mutdinov V International Conference "Petrochemistryl99': Nizhnekamsk, 1, 124-126 (1999). [loo] K. S. Minsker, N. R. Gil'mutdinov, A. D. Ishteryakov, R. Ya. Deberdeev, R. T. Shiyapov, G. S. Dyakonov, 0. V. Sofronova, R. G. Takhavutdinov, Al. Al. Berlin International Conference "Chemistry and Applied Chemistry - the Present and the Future", Sterlitamak, 22-24 (2000). [lo ll E. R. Islamov, Yu. A. Prochukhan, R. N. Gimaev Izvastiya VUZov. Khimiya i Khimicheskaya Tekhnologiya, 42,No.2, 73-76 (1999).

[lo21 K. S. Minsker, AI. Al. Berlin, V. P. Malinskaya, Ya. M. Abdrashitov, Yu. K. Dmitriev, V. D. Shapovalov, N. V. Zakharova International Conference "Chemistry and Applied Chemistry - the Present and the Future", Sterlitamak, 21-22 (2000). [lo31 M. G. Krekhova, S. K. Minsker, Yu. A. Prochukhan, K. S . Minsker Teoreticheskie Osnovy Khimicheskoi Technologii, 28,No. 3,271-273 (1994). [ 1041 A. F. Vurzel', A. L. Suris Izvastiya VUZov. Khimiya i Khimicheskaya Tekhnologiya, 40,No.

2, 116-118 (1997).

170

[lo51 M. G. Krekhova, K. S. Minsker, S. K. Minsker Teoreticheskie Osnovy Khimicheskoi Technologii, 29, No. 5,496-499 (1995). [lo61 E. V. Lebedeva, V. T. Sitenkov Khimiya i tekhnologiya topliv i masel, No. 1, 17-18 (1999). [lo71 V. F. Popov, N. V. Vinogradova Khimicheskaya Promyshlennost', No. 1,53-55 (1984). [lo81 V. F. Popov, N. V. Vinogradova Khimicheskaya Promyshlennost', No. 5,49-51 (1984). [lo91 A. A. Konoplev, G. G. Aleksanyan, B. L. Rytov, Al. Al. Berlin "Polimery-2002",IKhF RAN, MOSCOW, 24-26 (2002). [110] Al. Al. Berlin, K. S. Minsker, A. G. Mukhametzyanova, R. G. Takhavutdinov, G. S. D'yakonov, G. G. Aleksanyan, B. L. Rytov, A. A. Konoplev "Polimery-2002", IKhF RAN, MOSCOW, 121-127 (2003). [ill] A. A. Konoplev, G. G. Aleksanyan, B.L. Rytov, Al. Al. Berlin Teoreticheskie Osnovy

Khimicheskoi Technologii, 36, No. 2,220-222 (2002).

[112] V. M. Barabash, N. N. Smimov Zh. Prikl. Khimii, 67, No. 2, 196-203 (1994). [ 1131 A. M. Kutepov Chemical Dynamics and Theoretical Bases of Non-Linear Chemical-

Technological Processes, Moscow (1998) (in Russian).

[114] V. V. Konsetov, Yu. V. Kokotov Khimicheskaya Promyshlennost', No. 5,299-303 (1990). [ 1151 J. Baldyga, J. R. Bourne, R. V. Gholap Chemical Engineering Science, 50, No. 12, 1877-1880

(1995). [116] L. N. Braginskii, V. I. Begachev, V. M. Barabash Mixing in Liquid Mediums: Physical Bases and engineering calculation methods, Khimiya, Leningrad (1984) (in Russian). [ 1171 V. V. Bogdanov, E. I. Khristoforov, B. A. Klotsung Effective Low-Capacity Mixers, Khimiya,

Leningrad (1989) (in Russian). [I 181 N. S. Shulaev, I. Kh. Bikbulatov Bashkir. Khim. Zh., 4, No. 2,73-80 (1997). [ 1191 V. M. Barabash Teoreticheskie Osnovy Khimicheskoi Technologii, 28, No. 2, 110-117 (1994).

[ 1201 Yu. V. Svetlov Teoreticheskie Osnovy Khimicheskoi Technologii, 26, No. 6 , 819-835 (1992).

[121] A. N. Blaznov, V. A. Kunichan, D. V. Chashilov Zh. Prikl. Khimii, 74, No. 4, 621-624 (2001). [122] E. Lang, P. Ortina, F. Streiff, M. Fleischli International Journal Heat and Mass Transfer, 38, NO. 12,2239-2250 (1995). [123] H. Z. Li, C. Fasol, L. Choplin Chemical Engineering Science, 51, No. 10, 1947-1955 (1996). [124] D. M. Hobbs, F. J. Muzzio Chemical Engineering Science, 53, No. 18, 3199-3213 (1998).

171

[125] R. G. Takhavutdinov, G. S. D'yakonov, R. Ya. Deberdeev, K. S. Minsker Khimicheskaya Promyshlennost', No. 5,41-49 (2000). [ 1261 E. Turgeon, D. Pelletier, L. Ignat Paper of the 36 th Aerospace Sciences Meeting and Exhibit.

Proceeding of the American Institute of Aeronautics and Astronautics, 1-19 (1998).

[127] V. Z. Kompaniets, A. A. Konoplev, L. S. Polak in the book: Experimental and Theoretical Investigations of Plasmo-Chemical Processes, Khimiya, Moscow (1984) (in Russian). [ 1281 V. M. Olevskii, V. R. Ruchinskii Rotary-Membranous Heat- and Mass-Transfer Apparatus,

Khimiya, Moscow (1977) (in Russian). [ 1291 V. P. Budtov, V. V. Konsetov Heat- and Mass-Transmission in Polymerization Processes

Khimiya, Leningrad (1988) (in Russian). [130] J. Villermaux, L. Falk Chemical Engineering Science, 49,5127-5132 (1994). [131] P. Guichardon, L. Falk, J. Villermaux Chemical Engineering Science, 52, No. 24, 4649-4658 (1997). [132] J. Baldyga, J. R. Bourne, S. J. Hearn Chemical Engineering Science, 52, No. 24, 457-501 (1998). [133] J. Balduga, J. R. Bourne Chemical Engineering Journal, 42, 83-101 (1989). [ 1341 F. G. Galimzyanov, R. F. Galimzyanov The Theory of Internal Turbulent Flow,Ekspert, Ufa

(1999) (in Russian). [135] L. D. Landau, E. M. Lifshits Theoretical Physics. V. 6: Hydrodynamics, Nauka, Moscow (1986) (in Russian).. [136] V. A. Kaminskii, A. B. Rabinovich, A. Ya. Fedorov, V. A. Frost, A. A. Nur Zh. Fizicheskoi Khimii, 69, No. 8, 1456-1461 (1995). [137] V. A. Kaminskii, A. Ya. Fedorov, V. A. Frost. Teoreticheskie Osnovy Khirnicheskoi Technologii, 28, No. 6,591-599 (1994). [138] V. A. Kaminskii, A. B. Rabinovich, A. Ya. Fedorov, V. A. Frost Teoreticheskie Osnovy Khimicheskoi Technologii, 31, No. 3,243-249 (1997). [ 1391 A. Ya. Fedorov, V. A. Kaminskii Teoreticheskie Osnovy Khimicheskoi Technologii, 31, No.

2, 177-183 (1997). [140] R. I. Nigmatulin Dynamics of Multi-Phase Mediums, V. 1, Nauka, Moscow (1987) (in Russian).. [141] W. M. Long, L. V. Kalachev Chemical Engineering Science, 55,2295-2301 (2000). [ 1421 D.Maggioris, A. Goulas, A. H. Alexopoulos, E.G. Chatzi, C. Kiparissides Chemical

Engineering Science, 55,461 1-4627 (2000).

[143] A. A. Dontsov, G. Ya. Lozovik, S. P. Novitskaya Chlorinated Polymers, Khimiya, Moscow (1979) (in Russian)..

172

[144] N. M. Chirkov, P. E. Matkovskii, F. S. D'yachkovskii Polymerization on Complex Metalloorganic Catalysts, Khimiya, Moscow (1976) (in Russian).. [145] D. N. Latypov, A. A. Ovchinnikov Izvestiya VUZov. Khimiya i Khimicheskaya Tekhnologiya, 44, NO. 2, 108-110 (2001). [146] E. D. Shukin, A. V. Pertsov, E. A. Amelina Colloidal Chemistry, MGU, Moscow (1982) (in Russian). [ 1471 V. M. Barabash, M. A. Belevitskaya Teoreticheskie Osnovy Khimicheskoi Technologii, 29,

NO.4,362-372 (1995). [ 1481 M. A. Belevutskaya, V. M. Barabash Teoreticheskie Osnovy Khimicheskoi Technologii, 28,

NO.4, 342-347 (1994). [ 1491 W. Podgorska, J. Baldyga Chemical Engineering Science, 56,741-746 (2001). [ 1501 L. N. Braginskii, M. A. Belevitskaya Teoreticheskie Osnovy Khimicheskoi Technologii, 24,

NO.4, 509-513 (1990). [15 11 L. N. Braginskii, M. A. Belevitskaya Teoreticheskie Osnovy Khimicheskoi Technologii, 25, NO. 6, 843-852 (1991). [152] V. V. Yaroshenko, L. N. Braginskii, V. M. Barabash Teoreticheskie Osnovy Khimicheskoi Technologii, 22, No. 6,787-891 (1988). [ 1531 A. W. Patwardhan, J. B. Joshi Industrial Engineering Chemical Research, 36, 3904-3914

(1997). [154] V. N. Ivanets, S. N. Al'brekht, G. E. Ivanets Khimicheskaya Promyshlennost: No. 11, 46-48 (2000). [155] M. A. Promtov, A. I. Zimin Pharmaceutical Chemistry Journal, 34, No. 10,553-554 (2000). [156] Yu. V. Badikov, R. B. Valitov, V. S. Pilyugin Bashkir. Khim. Zh., 7, No. 2,58-68 (2000). [157] V. S. Pilyugin, R. B. Valitov, Yu. V. Badikov, G. E. Chekisheva, E. Sapozhnikov Bashkir. Khim. Zh., 7, No. 2, 69-75 (2000). [l58] V. S. Pilyugin, R. B. Valitov, Yu. V. Badikov, T. P. Vorob'eva, E. V. Klimakova, G. V. Kisileva, S. L. Kuznetsova, E. Sapozhnikov Bashkir. Khim. Zh., 7,No. 2,76-78 (2000). [159] M. Nadler, D. Mewes International Journal Multiphase Flow,33, No. 1, 55-68 (1997). [160] A. N. Kolmogorov Dokl. Akad. Nauk SSSR, 32, No. 1, 19-22 (1941). [161] B. E. Launder, D. B. Spalding Mathematical models of turbulence, Acad. Press 6 th, London, 16-35 (1972). [162] J. T. Kuo, G. B. Wallis International Journal Multiphase Flow,14, No. 5,547-571 (1988).

173

[ 1631 B. E. Launder, D. B. Spalding Computational methods in applied mechanics and engineering,

3,269-289 (1974). [164] Y. S. Chen, W. S. Kim NASA CR, 332-350 (1987). [165] V. Yakhot, S. A. Orsag Journal Science Computarion, 1,3-21 (1986) [166] V. Yakhot, S. A. Orsag, S. Thangam, T. B. Gatski, C. G. Speziale Physical Fluids, 4, No. 7, 14-25 (1992). [167] L. M. Smith, W. C. Reynolds Physical Fluids, 4, No. 2, 364-370 (1992). [168]. P. Bradshow, B. E. Launder, J. I. Lumley AZAA Paper, 97-120 (1991).

- European computational aerodynamics research project: validation of CFD codes and assessment of turbulence models. Notes on Numerical Fluid Mechanics, 58, 16-22 (1996).

[ 1691 TCARP

[ 1701 Documentation for PHOENICS. Cham Ltd. Wimbeldon, London (2000).

[171] D. B. Spalding Modelling of Turbulence. Imperial college CFDU report HTS, No. 2, 98-116 (1981). [172] W. Rodi Journal Geologycal Research, 92, No. 5,5305-5313 (1987). [ 1731 N. N. Kulov, E. K. Nicolaishvili, V. M. Barabash Chemical Engineering Communications,

21,259-266 (1983). [174] Yu. B. Monakov, A. G. Tolstikov Catalytic Polymerization of 1,3-dienes, Nauka, Moscow (1990) (in Russian). [ 1751 B. A. Dolgoplosk, E. I. Tinyakova Organometallic Catalysis in Polymerization Processes,

Nauka, Moscow (1985) (in Russian) [176] P. A. Kirpichnikov, V. V. Beresnev, L. M. Popova Album of Technological Schemes of the Base Productions in Synthetic Rubber Industry, Khimiya, Leningrad (1986) (in Russian). [ 1771 P. A. Kirpichnikov, L. A. Averko-Antonovich, Yu. 0. Averko-Antonovich Chemistry and

Technology of Synthetic Rubber, Khimiya, Leningrad (1975) (in Russian).

[178] S. R. Ivanova, R. F. Nafikova, V. P. Zakharov, K. S. Minsker Trudy Academii Nauk RB, Gilem, Ufa (2001) (in Russian). [179] W. S. Kim, J. M. Tarbel Chemical Engineering Communications, 176, 89-113 (1999). [180] J. Baldyga, W. Orciuch Power Technolog, 121,9-19 (2001). [181] M.-H. Sung, 1.4. Choi, J.-S. Kim, W.-S. Kim Chemical Engineering Science, 55, 2173-2184 (2000). [ 1821 J. Baldyga, W. Orciuch Chemical Engineering Science, 56,2435-2444 (2001).

174

[183] J. Baldyga, W. Podgorska, R. Pohorecki Chemical Engineering Science, 50, No. 8, 12811300 (1995). [184] J. Chen, C. Zheng, G. Chen Chemical Engineering Science, 51, No. 10, 1957-1966 (1996). [185]. N. V. Lisitsin Khimicheskaya Promyshlennost', 80, No. 1,47-53 (2003). [ 1861 G. S. D'yakonov, R. G. Takhavutdinov, I. V. Averko-Antonovich Teoreticheskie Osnovy

Khimicheskoi Tekhnologii, 35,No. 4,429-434 (2001). [ 1871 I. V. Averko-Antonovich, R. G. Takhavutdinov, G. S. D'yakonov Khimicheskaya

Promyshlennost', No. 5,52-56 (1999).

[188] V. A. Kirillov, N. A. Kuzin, A. B. Shigarov, M. M. Danilova, V. I. Drobyshevich Teoreticheskie Osnovy Khimicheskoi Tekhnologii, 32, No. 4,422-432 (1998). [189] G. W. Koning, K. R. Westerterp Chemical Engineering Science, 54,2527-2533 (1999). [190] Yu. G. Chesnokov Zh. Prikl. Khimii, 70,No.l,907-912 (1997). [191] B. V. Dzubenko, G. A. Dreitser, A. V. Kalyatka International Journal Heat and Mass Transfer, 41, No.3, 645-651 (1998). [192] B. V. Dzubenko, G. A. Dreitser, A. V. Kalyatka International Journal Heat and Mass Transfer, 41, No.3, 653-661 (1998). [ 1931 E. K. Kalinin, G. A. Dreitser Intensijkation of Heat-Exchange in Canals, Mashinostroenie,

Moscow (1990) (in Russian) [194] G. A. Dreitser Teploenergetika, No.3, 11-18 (1995). [195] Inventors Certificate 1223016 SSSR V. I. Gomon, P. G. Ostapushenko, G. A. Dreitser, E. K. Kalinin (1986). [196] Inventors Certificate 1307211 SSSR G. A. Dreitser, E. K. Kalinin, A. S . Neverov, N. V. Paramonov, E. S. Levin, S. G. Zakirov, K. V. Smol'skii (1987). [197] E. K. Kalinin, G. A. Dreitser, E. V. Dubrovsky Heat Transfer Engineering, 6 , No.1, 44-51 (1985). [198] G. A. Dreitser Modern Problems of Hydrodynamics and Geat-Exchange in Elements of Energy Devices and Cryogenic Technics, VZMI, Moscow (1988) (in Russian). [199] A. A. Konoplev, G. G. Aleksanyan, B. L. Rytov, Al. Al. Berlin "Polimery-2000", IKhF RAN, MOSCOW, 9-16 (2000). [200] Patent 21 11432 Russia A. A. Konoplev (1998). [201] F. Balzereit, W. Roetzel Chemical Ing. Technology, 69, No.9,2324-2330 (1997). [202] A. A. Gukhman Introduction into Similarity Theory, Vysshaya Shkola, Moscow (1963) (in Russian)

175

[203] M. Zlokarnik Chemical Engineering Science, 53,3023-3030 (1998). [204] 0. Florea, 0. Smigel'skii Calculations of Processes and Apparatus of Applied Chemistry, Khimiya, Moscow (1977) (in Russian) [205] R. Kh. Rakhimov, P. I. Kutuzov, Yu. P. Bazhenov, I. Sh. Nasyrov Bashkir. Khim. Zh., 4, No.2, 14-17 (1997). [206] I. R. Mullagaliev, Yu. B. Monakov Vysokomol. Soedin., 44, No.12,2251-2292 (2002). [207] I. V. Garmonov Synthetic Rubber, Khimiya, Leningrad (1983) (in Russian). [208] G. A. Aminova, G. V. Manuiko, G. I. Litvinenko, G. S. Dyakonov, V. L. Zolotarev, A. G. Filipva, T. V. Ignashina, V. V. Davydova, 0.V. AntonovaZh. PrikL Khimii, 75, No.7,1169-1173 (2002). [209] K. S. Minsker, M. M. Karpasas, G. E. Zaikov Uspekhi Khimii, 55, No.l,29-61 (1986). [210] K. S. Minsker, 0. A. Ponomarev, M. M. Karpasas, G. S. Lomakin, Yu. B. Monakov Vysokomol. Soedin., 24A, No.7, 1360-1367 (1982). [211] K. S. Minsker, M. M. Karpasas, Yu. B. Monakov, I. R. Mullagaliev Vysokomol. Soedin., 29B, No.8, 608-611 (1987). [2121 K. S. Minsker, M. M. Karpasas, Yu. B. Monakov, 0. A. Ponomarev, S. R. Rafikov Dokl. Akad. Nauk SSSR, 266, No.3, 658-661 (1982). [213] K. S. Minsker, M. M. Karpasas Teoreticheskaya i Eksperimental'naya Khimiya, 22, No.2, 160-168 (1986). [214] N. N. Sigaeva, V. G. Kozlov, Kh. K. Gadeleva, V. K. Nefedev, Yu. B. Monakov Bashkir. Khim. Zh., 5, No.2.27-30 (1998). [215] N. N. Sigaeva, V. G. Kozlov, K. V. Nefed'ev, N. G. Marina, Yu. B. Monakov Bashkir. Khim. Zh., 5, No.2, 30-34 (1998). [216] H. Iovu, G. Hubca, D. Racoti, J. S. Hurst European Polymer Journal, 35, 335-344 (1999). [2 171 S. Ya. Frenkel' Introduction into Statistical Theory of Polymerization Nauka, Moscow, Leningrad (1965) (in Russian). [218] V. P. Budtov, E. G. Zotikov, E. L. Ponomareva, M. I. Gandel'sman Vysokomol. Soedin., 27A, NO.5, 1094-1097 (1985). [219] Yu. B. Monakov, N. N. Sigaeva Vysokomol. Soedin., 43C, No.9, 1667-1688 (2001). [2203 N. N. Sigaeva, T. S. Usmanov, E. A. Shirokova, V. P. Budtov, S. I. Spivak, Yu. B. Monakov Dokl. Akad. Nauk SSSR, 365, No.2, 221-224 (1999). [221] Sigaeva N.N., Usmanov T.S., V. P. Budtov, S. I. Spivak, Yu. B. Monakov Vysokomol. Soedin., 42B, No.1, 112-117 (2000).

176

[222] N. N. Sigaeva, E. A. Shirokova, I. R. Mullagaliev, I. A. Ionova, V. P. Budtov, Yu. B. Monakov Vysokomol. Soedin., 42A, No.8, 1269-1274 (2000). [223] N. N. Sigaeva, T. S. Usmanov, V. P. Budtov, S. I. Spivak, Yu. B. Monakov Zh. Prikl. Khimii, 7 4 , No.7, 1110-1116 (2001). [224] E. I. Vizen, F. I. Yakobson Vysokotnol. Soedin., 20A, No.4,927-935 (1978). [225] A. N. Tikhonov, V. Ya. Arsenin Methods of Solution of Ill-Posed Problems Nauka, Moscow (1986) (in Russian). [226] A. N. Tikhonov, A. V. Goncharskii, V. V. Stepanov, A. G. Yagola Numerical Methods IllPosed Problems Solution, Nauka, Moscow (1990) (in Russian). [227] G. Frolov Colloid Chemistry. Su$ace phenomena and Disperse systems, Khimiya, Moscow (1989) (in Russian). [228] E. Magni, G.A. Somorjai Sugace Science, 377-379, 824-827 (1997). [229] V. A. Grechanovskii, L. G. Andrianova, L. V. Agibalova, A. S. Estrin, I. Ya. Poddubnyi Vysokomol. Soedin., 22A, No.9,2112-2120 (1980). [230] Yu. A. Gavrilov, A. M. Aladoshev, N. Kovaleva, L. A. Novoshenova, N. S. Bukharkina, V. P. Konovalov Vysokomol. Soedin., 27A, No.ll,2300-2304 (1985). [231] V. S. Bodrova, E. P. Piskareva, L. F. Shelokhneva, I. L. Poletaeva Vysokomol. Soedin., 40A, No.5, 1741-1749 (1998). [232] I. A. Jaber, K. Hauschild, G. Fink Macromolecul Chemistry, 191, No.9, 2067-2076 (1990). [233] U. S. Agarwal Chemical Engineering Science, 53,3941-3949, (1998). [234] V. N. Karasev, K. S. Minsker Vysokomol. Soedin., 15A, No.5,955-960 (1973). [235] V. N. Karasev, Al. Al. Berlin, K. S. Minsker Vysokomol. Soedin.,l6A, No.6, 1300-1305 (1974). [236] G. S. D'yakonov, R. G. Takhavutdinov, L. M. Kurochkin, A. G. Mukhametzyanova Vestnik Kazanskogo Tekhnologicheskogo Universiteta, No. 1, 70-75 (1998). [237] G. S. D'yakonov, R. G. Takhavutdinov, A. G. Mukhametzyanova Mass-Exchange Processes and Apparatus of Applied Chemistry, KGTU, Kazan (1998) (in Russian). [238] S. A. Buder, G. A. Perlin, V. V. Solodskii, L. S. Egova Promyshlennos'r SK, No.7,9-1 (1983). [239] Z. M. Borovkova, L. V. Kosmodem'yanskii, E. P. Kopylov Poluchenie i Svoistva Polimerov nu osnove piperilena, CNIITENeftekhim, Moscow (1983) (in Russian). [240] F. N. Kaputskii, V. P. Mardykin Kationnyii Oligomer Piperilena: Sintez, Svoistva, Primenenie, BGU, Minsk (1997) (in Russian).

177

[241] R. M. Livshits, L. A. Dobrovinskii Zameniteli Rastitel'nykh Masel v Lakokrasochnoi Provyshlennosti, Khimiya, Moscow (1987) (in Russian). [242] N. S. Enikolopyan, E. F. Oleinik, Entsiklopediya Polimerov, Sovetskaya Entsiklopediya, Moscow, 1,973 (1972). [243] J. P. Kennedy, E. Marechal Carbocationic polymerization, Wiley, New-York (1982). [244] T. T. Denisova, I. A. Livshits, E. R. Gershtein Vysokomol. Soedin., 16A, No.4, 880-885 (1974). [245] N. N. Sigaeva, V. G. Kozlov, K. V. Nefed'ev, 0. I. Kozlova, T. S. Usmanov, E. A. Shirokova, Yu. B. Monakov Bashkir. Khim. Zh., 5 , No.5, 31-36 (1998). [246] B. L. Irkhin, V. I. Ponomarenko, K. S. Minsker Promyshlennost: SK, No.9, 18-20 (1974). [247] S. A. Egoricheva, V. A. Rozentsvet, B. I. Pantukh, R. M. Livshits Promyshlennost' SK,No.11, 7-12 (1985). [248] V. A. Rozentsvet, S. A. Egoricheva, Zh. A. Matveeva, B. I. Pantukh, R. M. Livshits Promyshlennost' SK, No.4, 8-12 (1987). [249] V. A. Babkin, G. E. Zaikov, K. S. Minsker Quantum-Chemical Aspects of Cationic Polymerization of Olefines, Gilem, Ufa (1996) (in Russian). [250] V. P. Mardykin, S. G. Morozova, L. V. Gaponik, L. D. Chuprakova Zh. Prikl. Khimii, 71, No.6, 1041-1043 (1998). [251] Patent 2044742 Russia Yu. N. Zhidkov, M. P. Sedach, V. I. Potkin, A. A. Erdman, I. A. Yursha, L. A. Olenkevich, L. M. Bondarenko (1995). [252] V. P. Malinskaya, V. P. Zakharov, K. S. Minsker, R. Kh. Rakhimov, P. I. Kutuzov, Z. A. Abzalin, L. M. Kurochkin, R. Ya. Deberdeev The Vth International Conference "Nefekhimiya99", Nizhnekamsk, 1,78-79 (1999). [253] T. G. Burganov, T. G. Vaseneva, I. V. Gladkov, N. K. Zaripova, S. E. Zhukova, N. M. Evtishina, L. M. Kurochkin, V. A. Mikheeva, B. R. Serebryakov, V. V. Samoilova The Vth International Conference "Neftekhimiya-99", Nizhnekamsk, 1,88-89 (1999). [254] Patent 2141871 Russia R. Ya. Deberdeev, K. S. Minsker, Al. Al. Berlin, G. S. D'yakonov, E. I. Nagumanova, L. M. Kurochkin, Kh. V. Mustafin, N. R. Gil'mutdinov, Yu. I. Ryazanov, V. A. Shamanskii, V. P. Pogrebtsov, A. I. Vorob'ev, T. G. Burganov, R. G. Salakhutdinov, N. P. Boreiko (1999). [255] Patent 2141872 Russia Al. AI. Berlin, K. S. Minsker, R. Ya. Deberdeev, R. G. Galiev, Yu. I. Ryazanov, A. Sh. Ziyatdinov, V. P. Pogrebtsov, Z. A. Abzalin, T. G. Burganov, A. I. Vorob'ev, A. A. Blinov, G. V. Baev (1999). [256] V. A. Myagchenkov, S. Ya. Frenkel' Compositional Heterogeneity of Copolymers, Khimiya, Leningrad (1988) (in Russian).

178

[257] A. G. Mukhametzyanova, G. Z. Sakhapov, G. S. D'yakonov, R. G. Takhavutdinov, L. M. Kurochkin Mass-Exchange Processes and apparatus of Applied Chemistry, KGTU, Kazan (1995). [258] G. S. D'yakonov, R. G. Takhavutdinov, L. M. Kurochkin, G. Z. Sakhapov The IV Conference on intensification of petrochemical processes, Nizhnekamsk (1996). [259] A. A. Konoplev, G. G. Aleksanyan, B. L. Rytov, AI. Al. Berlin Polimery-2001, IKhF RAS, Moscow (2001). [260] R. M. Khusainova, V. P. Dorozhkin, D. A. Maksimov The Vth International Conference "Nefi'ekhimiya-99",Nizhnekamsk, 2, 248 (1999). [261]. K. S. Minsker, AI. AI. Berlin, R. Ya. Deberdeev, S. R. Ivanova Khimicheskaya Promyshlennost', No.11, 26-30 (2000). [262] V. S. Shmarlin Synthesis, properties and application rubbers,CNIITENeftekhim, Moscow (1973) (in Russian).

of

modiped

butyl

[263] K. S. Minsker, R. F. Gataullin, V. V. Lisitskii, I. S. Beresneva, S. B. Sal'nikov, V. S. Shmarlin Vysokomol. Soedin., 25A, No.8, 1686-1690 (1983). [264] K. S. Minsker, R. F. Gataullin, V. M. Yanborisov, T. A. Krasnova, S. B. Sal'nikov, V. S. Shmarlin Vysokomol. Soedin., 26A, No.4,781-784 (1984). [265] Ya. N. Prokof'ev, N. V. Sherbakova, V. P. Bugrov Synthetic Rubber, Khimiya, Leningrad, 342-355 (1976) (in Russian).. [266] N. V. Abramova, L. P. Bataeva, E. M. Sire, Yu. N. Orlov, S. P. Korshunov Khimicheskaya Promyshlennost', No.12, 1-3 (2001). [267] L. A. Oshin Industrial Chloro-Organic Products, Khimiya, Moscow (1978) (in Russian).. [268] I. Dekhant Infa-red SpectroscopyOfPolymers,Khimiya, Moscow (1976) (in Russian). [269] S. R. Rafikov, V. P. Budtov, Yu. B. Monakov Introduction into Physico-Chemistry of Polymers Solutions, Nauka, Moscow (1978) (in Russian). [270] V. I. Kimel'blat, M. G. Khakimov, I. G. Chebotareva, S. I. Vol'fson Mekhanika Kompozitsionnykh Materialov, 34, No.4,531-536 (1998). [271] V. I. Kimel'blat, S. I. Vol'fson, I. G. Chebotareva, T. V. Malysheva Mekhanika Kotnpozitsionnykh Materialov, 34,No.5, 691-695 (1998). [272] I. V. Garmonov Analysis of Products Synthetic Rubbers Production, Khimiya, Moscow (1964) (in Russian). [273] Yu. G. Frolov, A. S. Grodskii Laboratory Works and Problems of Colloid Chemistry, Khimiya, Moscow (1986) (in Russian). [274] 0. Zenkevich, K. Morgan Terminal Elements and Approximation, Mir, Moscow (1986) (in Russian).

179

[275] M. C. Chaturvedi Journal of the hydraulics division. Proceedings of the American Society of civil Engineers, 89,61-92 (1963). [276] S. M. Loktev, V. L. Klimenko, V. V. Kamozoskin, A. T. Menyailo The Higher Butters, Khimiya, Moscow (1970) (in Russian). [277] B. R. Serebryakov, T. K. Plaskunov, V. R. Nizheles, M. A. Dalin The Highest Olefnes, Khimiya, Moscow (1984) (in Russian).