Thermal Separation Processes

Klaus Sattler, Hans Jacob Feindt

Thermal Separation Processes

0 VCH Vcrlagsgescllschaft mbH, D-60451 Wcinheim (Federal Rcpuhlic ot Germany). IWS

Distribution:

VCH, P.0. Box 10 1161. D-69451 Weinheim. Federal Kcpublic of Germany

Switzerland: VCH. P.O. Box. CH-4020 Bascl. Switzerland

LJnitcd Kingdom and Ireland: VCH, 8 Wellington Court. Cambridge CBI 1HZ. United Kingdom USA and Canada: VCH, 220 East 23rd Street. New York, NY 10010-4606. USA

Japan: VCH. Eikuw Building, 10-9 Hongo 1-chomc. Bunkya-ku, Tokyo 113. Japan

ISBN 3-527-28622-5

Klaus Sattler, Hans Jacob Feindt

Thermal Separation Processes Principles and Design

Weinheim - New York Base1 - Cambridge - Tokyo

Prof. Dipl.-Ing. Klaus Sattler Fachhochschule fur Technik Speyerer Stral3e 4 D-68163 Mannheim

Dr. Hans Jacob Feindt BASF AG Abteilung Verfahrenstechnik D-67056 Ludwigshafen

This book was carefullyproduced. Nevertheless, authors and publisher donotwarrant the information contained therein to be free of errors. Readers are advised to keep in mind that statements, data, illustrations, procedural details or other items may inadvertently be inaccurate.

Published jointly by VCH Verlagsgesellschaft. Weinheim (Federal Republic of Germany) VCH Publishers. New York. NY (USA)

Editorial Directors: Philomena Ryan-Bugler, Louise Elsam, Karin Sora Production Manager: Claudia Gross1

Library of Congress Card No. applied for A catalogue record for this book is available from the British Library

Die Deutsche Bibliothek - CIP-Einheitsaufnahme Sattler, Klaus: Thermal separation processes : principles and design / Klaus Sattlcr ; Hans Jacob Feindt. - 1. ed. Wcinhcim ; Ncw York ; Bascl ; Cambridge ; Tokyo : VCH, 1995 ISBN 3-527-28622-5 (Weinheim ...) N E : Feindt, Hans Jacob:

0VCH Verlagsgesellschaft m b H , D-69451 Wcinhcim (Federal Rcpublie of Germany), 1995 Printed on acid-free and low-chlorine paper

All rights reserved (including those of translation into other languages). No part of this book may be reproduced in any form-byphotoprinting,microfilm.or anyother means-nortransmittedartranslated intoamachinelanguage without written permission from the publishers. Registered names, trademarks, etc. used in this book. even when not specifically marked as such, are not to bc considcrcd unprotected by law. Composition: Filmsatz Unger & Sommcr GmbH, D-69469 Weinheim Printing and Bookbinding: Druck haus ,Thomas Muntzer" GnibH, D-99947 Bad Langcnsalza Printed in the Federal Repuhlie of Germany

Foreword

The separation of gaseous and liquid solutions into their components and the drying of wet products have always been an integral part of the manufacture of products in the chemical, petroleum, food, and pharmaceutical industries. As environmental protection has become an increasingly important consideration to industry, separation processes have become more important in direct proportion. This book provides a clear fundamental development of the technology of important separation processes. As indicated by the title the book deals with separation processes in which heat is an input to the complete process of separating the constituents of a mixture. The flow of heat in the process is clear in distillation, crystallization and drying but is not so obvious in absorption, extraction and adsorption, where the heat flow is required to regenerate the solvent or adsorbent.

Each of these six subjects is given thorough coverage in its own chapter. These chapters follow a comprehensive development of the physical chemistry and engineering which provide the principles upon which the separation processes are based. The individual process treatments cover computational algorithms, equipment design criteria and energy conservation. The overall treatment permits the evaluation of competing separations techniques and the choice of the optimal process. This book is intended as a college or university level text for students in chemical engineering and related fields. It is also complete enough and detailed enough in its development of each topic to be useful as a reference for practicing engineers both new to and experienced in the area of separations. CCNY, New York May 1994

Prof. H. Weinstein

Preface

This book, transformed from the original German by Dr. H. J. Feindt, is based on two German editions “Thermische Trennverfahren”, published by Prof. K. Sattler. They have been successfully used as textbooks for university and college students and as reference texts in seminars and training programs for practising engineers in Germany, Austria and Switzerland. The book presents a clear and very practice-oriented overview of thermal separation technologies. An extensive introduction elucidates the physical, physico-chemical, and chemical engineering fundamentals and principles of the different unit operations used to separate homogenous gaseous and liquid mixtures. The introduction is followed by a concise text with many explanatory figures and tables referring to process and basic design, flow-sheets, basic engineering and examples for the application of the unit operations distillation, absorption, adsorption, drying, liquid-liquid and solid-liquid extraction, evaporation and crystallization of solutions, melt crystallization and desublimation. A comprehensive reference list allows follow up of special separation problems. The book enables the reader to choose and evaluate thermal separation processes and to model and design the necessary separation plant equipment.

Chemical and mechanical engineers, chemists, physisists, bio-technologists in research and development, plant design, production, environmental protection and administration and students in engineering and natural sciences will find this treatment of exceptional value and practical use. Due to the quantity of the topics covered exercises could not be included in this book. An additional collection of illustrations with reference to basic engineering and design of the necessary equipment of thermal separation units is available in German (Sattler, K. : Thermische Trennverfahren. Aufgaben und Liisungen, Auslegungsbeispiele) and will be translated into the English language. We are very much obliged to Prof. H. Weinstein, City University of New York for his advice and his Foreword to this book. Many thanks are also given to Philomena Ryan-Bugler, Louise Elsam, Karin Sora and the production team of VCH Verlagsgesellschaft for the accurate lectorship and book production. Special thanks are also given to Paul Fursey, University of Bradford, United Kingdom, for his assistance in copy-editing. Briihl, Ludwigshafen December 1994

K. Sattler H. J. Feindt

Contents

Frequently Used Nomenclature XV

1

Basic Concepts 1

1.1

Principles of Thermal Separation Processes

1.2

Thermal Separation Process Modes

1.3 1.3.1 1.3.2 1.3.3

Mass Balance, Energy Balance, Exergy Balance Mass, Energy and Heat Balances 8 Exergy Balance 12 Calculation of Balance Equations 13

1.4 1.4.1 1.4.1.1

1.4.6

Phase Equilibria 14 Basic Concepts 14 General Differential Equation for the Equilibrium Between Two Phases 17 The Gibbs Phase Rule 18 Liquid-Liquid Equilibrium 19 The Nernst Distribution Law 19 Representation of Liquid-Liquid Phase Equilibrium 23 Vapor-Liquid Equilibrium 28 One Component Systems 28 Two and Multicomponent Systems 30 Henry’s Law, Gas Solubility 44 Boiling Equilibrium of a Solid Solution, Decrease of Vapor Pressure and Increase of Boiling Point 51 Gas-Solid Phase Equilibrium 52 Gas-Solid Phase Equilibrium, Sublimation 52 Gas-Solid Phase Equilibrium with Adsorption/Desorption and Convective Solid Drying (Adsorption Equilibrium) 54 Liquid-Solid Phase Equilibrium 60 Solubility of Solids in Liquid Solvents 60 Melting Pressure Curve 62 Decrease in the Freezing Point 63 State Diagrams of Binary Systems for Solid and Liquid Phase Equilibrium 65 Enthalpy of Phase Changes 65

1.5

Separation Factor and Relative Volatility

1.6

Minimum Separation Work

1.4.1.2 1.4.2 1.4.2.1 1.4.2.2 1.4.3 1.4.3.1 1.4.3.2 1.4.3.3 1.4.3.4 1.4.4 1.4.4.1 1.4.4.2 1.4.5 1.4.5.1 1.4.5.2 1.4.5.3 1.4.5.4

67

1

7

67

8

X

Contents

1.7 1.7.1 1.7.1.1 1.7.1.2 1.7.1.3 1.7.2 1.7.3 1.7.3.1

Mass Transfer Fundamentals 68 Mass Transfer by Molecular Diffusion 69 Steady-State Diffusion 69 Unsteady-State Diffusion 70 Diffusion Coefficient 70 Mass Transfer by Convection 72 Overall Mass Transfer 74 Two Film Theory, Mass Transfer Coefficient and Turbulence Theory

1.8

Steady-State Cocurrent Operation

1.9 1.9.1 1.9.2

1.9.4 1.9.4.1 1.9.4.2

Steady-State Countercurrent Operation 79 Theory of Separation Stages 79 Method to Determine the Number of Theoretical Separation Stages for a Countercurrent Column 82 Calculation for Counterflow Columns 86 Mass Balances 89 Phase Equilibrium Relationship 89 Enthalpy Balances 89 Stoichiometric Conditions for the Sum of the Concentration at Each Equilibrium Stage 90 Kinetic Theory for the Counterflow Separation of a Mixture 90 Two-Directional Mass Transfer Between Phases 91 One-Directional Mass Transfer 92

1.10

Steady-State Crossflow Operation

1.11

General Procedure to Design Equipment for the Thermal Separation of Mixtures 94

2

Distillation and Partial Condensation 101

2.1

Concepts of Simple Distillation, Rectification and Partial Condensation 101

2.2 2.2.1 2.2.2 2.2.3 2.2.4

Discontinuously and Continuously Operated Simple Distillation, Flash Distillation 103 Discontinuously Operated Simple Distillation 103 Continuously Operated Simple Distillation 107 Heat Requirement of Simple Distillation Units 109 Flash Distillation 111

2.3

Carrier Distillation

2.4

Vacuum and Molecular Distillation

2.5 2.5.1 2.5.1.1

Countercurrent Distillation (Rectification) 119 Process Variations of Rectification 119 Continuously Operated Rectification in Rectification Columns with Enriching and Stripping Zones 119 Stripping (Exhausting) Column 120

1.9.3 1.9.3.1 1.9.3.2 1.9.3.3 1.9.3.4

2.5.1.2

75

77

94

113 116

Contents

XI

2.5.4 2.5.5 2.5.6 2.5.6.1 2.5.6.2

Enrichment Column 121 Carrier Rectification 123 Combinations of Different Variations 123 Rectification with an Entrainer 123 Heteroazeotropic Rectification 129 Two Pressure Operation 130 Diffusion Distillation 131 Overpressure, Low Temperature and Vacuum Rectification 132 Continuous Adiabatic Rectification 134 Flow Rates 135 Heat Requirement of a Column 136 Energy Saving Steps 138 Determination of the Number of Separation Stages and Column Height for Heat and Mass Transfer 147 Minimum Reflux Ratio, Optimal Economic Reflux Ratio 157 Feed Stage 157 Discontinuous Adiabatic Rectification 158 Amount of Overhead Product 160 Heat Requirement 161 Still Diameter, Free Vapor Space, Column Diameter 162 McCabe-Thiele Method to Determine the Number of Theoretical Separation Stages 163 Semicontinuous Adiabatic Rectification 163 Determination of the Column Diameter 164 Internals in Rectification Columns 165 Column Trays 167 Random Packing, Packing with Regular Geometry 196

2.6

Choice, Optimization and Control of Rectification Units

2.7

Rectification Units Accessories 218

2.8

Parallel Flow Distillation

2.9

Nonadiabatic Rectification

2.10

Partial Condensation

3

Absorption

3.1 3.1.1 3.1.2

Principle of Absorption and Desorption, Processes and Process Examples 239 Concepts and Process Examples 239 Process Examples 240

3.2

Requirements of the Wash Liquid or Solvent, Solvent Consumption 243

3.3

Enthalpy and Heat Balances

3.4

Cocurrent Phase Flow Absorption

2.5.1.3 2.5.1.4 2.5.1.5 2.5.1.6 2.5.1.7 2.5.1.8 2.5.1.9 2.5,l. 10 2.5.2 2.5.2.1 2.5.2.2 2.5.2.3 2.5.2.4 2.5.2.5 2.5.2.6 2.5.3 2.5.3.1 2.5.3.2 2.5.3.3 2.5.3.4

216

222 222

230

239

246 248

XI1 3.5 3.5.1 3.5.2

Contents

Countercurrent Phase Flow Absorption, Design of Countercurrent Flow Columns 248 Determination of the Column Cross-Sectional Area 248 Determination of the Number of Stages and Column Height for Mass and Heat Transfer 250

3.6

Types of Absorber

3.7

Regeneration of the Solvent, Desorption

4

Adsorption 281

4.1 4.1.1 4.1.2

Principles of Adsorption and Desorption, Processes and Examples 28 1 Concept 281 Processes and Examples 282

4.2 4.2.1 4.2.2 4.2.3

Adsorbents, Selection of Adsorbent 291 Adsorbents 291 Requirements for the Adsorbent, Adsorbent Selection 291 Technical Adsorbents, Characteristic Data of Adsorbents 293

4.3

Adsorption Kinetics

4.4 4.4.1 4.4.2 4.4.3

Variation of Adsorption, Design of Adsorbers 301 Single Stage Adsorption in a Vessel Adsorber with Adsorbent Packing Multistage Adsorption with Cross Flow of Gas and Adsorbent Phases 307 Multistage Countercurrent Adsorption 308

4.5

Adsorber Types 310

4.6

Desorption, Regeneration of Loaded Adsorbent

5

Drying

5.1

Concepts, Processes and Examples

5.2

Characteristics of the Moist Product, Movement of Moisture

5.3

Properties of Wet Gases, h-X Diagram

5.4

Mass and Heat Transfer in Convection Drying

5.5

Drying Kinetics, Course of Drying, Drying Time

5.6 5.6.1 5.6.2 5.6.3

Convection Drying 340 Drying Gas and Heat Requirements in Convection Drying Steps in Energy Saving 343 Variations of Convection Drying 346

5.1

Contact Drying

5.8

Radiation Drying

262 263

293

311

317

349 351

317 320

324 331 335 340

301

Contents

5.9

Dielectric Drying

5.10

Freeze Drying (Sublimation Drying) 355

XI11

352

Design of Dryers 357 Overview of Dryers, Dryer Selection and Design 357 Individual Presentation of Selected Dryer Types with Design Aids 363 Chamber Dryer 363 Tunnel Dryer 364 Belt Dryer 364 Multiple Plate Dryer 364 Rotary Dryer 364 Fluidized Bed Dryer 366 Air-Flow Dryer, Pneumatic (Flash) Dryer 374 Spray Dryer 377 Drum Dryer 381 Thin Film Evaporation Dryer (Vertical and Horizontal Dryer) 381 5.1 1.2.11 Contact-Mixing Dryer 381 5.11.2.12 Contact Dryer with Continuous Product Movement due to Gravity 385 5.11.3 Process Control of Dryers 387 5.11 5.11.1 5.11.2 5.1 1.2.1 5.11.2.2 5.1 1.2.3 5. 11.2.4 5.11.2.5 5.11.2.6 5.11.2.7 5.1 1.2.8 5.11.2.9 5.1 1.2.10

6

Extraction 393

6.1

Basic Concepts and Processes 393

6.2 6.2.1 6.2.2 6.2.3 6.2.3.1 6.2.3.2 6.2.3.3 6.2.3.4 6.2.3.5 6.2.3.6 6.2.4 6.2.4.1 6.2.4.2 6.2.5

Liquid-Liquid Extraction 395 Fields of Application and Process Examples 395 Solvent Requirements, Selection of Solvent 399 Liquid-Liquid Extraction Variations 400 Single Stage Extraction 400 Differential Stagewise Extraction 403 Multistage Cross-Current Extraction 403 Multistage Countercurrent Extraction 407 Countercurrent Extraction with Extract Reflux 421 Countercurrent Distribution 424 Design Forms of Extraction Apparatus 424 Mixer-Settler, Mixer-Settler Cascade 425 Countercurrent Columns with and without Energy Supply 426 Selection and Design of Extraction Apparatus 456

6.3

Solid-Liquid Extraction (Leaching) 458

6.4

High Pressure Extraction (Distraction) 463

XIV

Contents

7

Solvent Evaporation, Crystallization 475

7.1

Basic Concept and Processing Modes of Crystallization

7.2 7.2.1 7.2.1.1 7.2.1.2 7.2.1.3 7.2.1.4 7.2.1.5 7.2.2 7.2.2.1 7.2.2.2 7.2.3 7.2.4 7.2.5 7.2.6 7.2.7

Crystallization from a Solution 484 Concentration of Solutions by Evaporation 485 Single Stage Solution Evaporation 486 Multistage Solution Evaporation 487 Solution Evaporation with Mechanical and Thermal Vapor Compression 492 Multistage Flash Evaporation 498 Types of Evaporators to Concentrate Solutions 500 Balancing of Crystallizers 500 Crystal Product Rate 500 Heat Exchange During Crystallization 506 Crystallization Kinetics, Crystal Seed Formation, Crystal Growth 508 Design of Crystallizers for Mass Crystallization from a Solution 511 Criteria for the Selection and Design of Crystallizers 516 Freezing 519 Fractional Crystallization of a Solution 520

7.3

Crystallization from a Melt

7.4

Crystallization from a Vapor Phase, Sublimation and Desublimation

8

Documentation and Calculation of Physical Characteristics 533

General References 537 Index 539

475

521 524

Frequently Used Nomenclature

A

Area

m2

AQ

Cross sectional area, cross section

m2

D D,D

Diffusion coefficient

m2/h

Vapor; vapor flow rate

kg, kmol; kg/h, kmol/h

E

Enrichment ratio, stage efficiency factor

F

Force

N

F

Loading factor for column trays

m/s

fiF

Feed; feed flow rate

kg, kmol; kg/h, kmol/h

F

Free internal energy

kJ

G, G

Gas; gas flow rate

kg, kmol; kg/h, kmol/h

G

Free enthalpy, Gibbs free energy

kJ

H

Enthalpy

kJ

HE TS

Height equivalent to one theoretical stage

m

HTU

Height of a transfer unit

m

K*

Phase equilibrium constant, distribution coefficient

L, L

Liquid; liquid flow rate

kg, kmol; kg/h, kmol/h

Lc M

Characteristic length

m

Molar mass

kg/kmol

N

Number of stages

NTU

Number of transfer units

Q,Q

Heat; heat flow rate

kJ; kJ/h, W

R, R

Reflux; reflux flow rate

kg, kmol; kg/h, kmol/h

R

Gas constant

kJ/(kmol K)

S

Entropy

k J/K

.

v m

+

=

1/pa

XVI

Frequently Used Nomenclature

T

Absolute temperature

K

U

Internal energy

kJ

V, V

Volume; volumetric flow rate

m3; m3/h

Molar volume

m3/kmol

Work

kJ

Ratio or loading of key component in liquid or heavy phase (moles i/moles inert, kg i/kg carrier (inert))

-

v

W

x

Y

Ratio or loading of key component in vapor or light phase (moles i/moles inert, kg i/kg carrier (inert))

Z

Length or height for heat and mass transfer

a

Activity

a

Specific volumetric area

m2/m3

cp, Cp

Specific heat

kJ/(kg. K), kJ/(kmol. K)

c

Molar concentration

kmol/m3

cw

Resistance coefficient

d

Diameter

m

dP

Particle diameter

m

ds

Sauter diameter

m

f

Fugacity

bar

Specific free internal energy

kJ/kg, kJ/kmol

gravitational acceleration

m/s2

Specific free enthalpy

kJ/kg, kJ/kmol

h,

Specific enthalpy

kJ/kg, kJ/kmol

Ah, A&

Latent heat

kJ/kg, kJ/kmol

k

Overall heat transfer coefficient

W/(m2. K)

k

Overall mass transfer coefficient

m/h

rn, m

Mass; mass flow rate

kg; kg/h

n, n

Number of moles; molar flux

kmol; kmol/h

P

Total pressure

bar, Pa

P,

Partial pressure of component i

bar, Pa

.Lf g g9

E

m

Frequently Used Nomenclature

XVII

Po, i

Saturated vapor pressure of component i

bar, Pa

AP

Pressure drop

mbar, Pa

4

Specific heat requirement

kJ/kg

4

Specific heat flux

kJ/(m2 h), W/m2

r

Radius

m

r

Reaction rate

kmol/(m3 h)

s, s

Specific entropy

kJ/(kg. K), kJ/(kmol K)

S

Characteristic distance (transfer distance)

m

t

Time

h

t,n

Mean residence time

h

u, ii

Specific internal energy

k J/kg, k J/kmol

W

Velocity

m/s

W

Mass fraction, weight fraction of component i

X

Molar fraction, heavy phase

Y

Molar fraction, light phase

Z

Variable distance length or height

Az

Tray spacing

-

m

Greek a

Separation factor

-

a

Heat transfer coefficient

W/(m2. K)

Mass transfer coefficient

m/h

Activity coefficient Film thickness, layer thickness

m

Porosity, void fraction of a bed of solids, fraction of free volume Yield

-

Dynamic viscosity

Pa s

-

Temperature

"C

Slope, gradient angle, inclination

0

Thermal conductivity

W/(m . K)

Chemical potential

k J/kg, kJ/kmol

XVIII

Frequently Used Nomenclature

V

Reflux ratio, solvent ratio, adsorbent ratio

-

V

Kinematic viscosity

m2/s

e

Density

kg/m3

0

Surface tension

N/m

v?

Relative humidity

-

Subscripts

H

Steam

T

Carrier

g is J

Gas phase Component

1

Liquid phase

0

Above, surface

P

Effective, practical

S

Solid phase

t

Theoretical

U

Below

U

Loss

a

Start, entry

0

End, exit

Thermal Separation Processes: Principles and Design Klaus Sattler, Hans Jacob Feindt copyright 0VCH

Verlagsgesellschaft m h H , 1995

Basic Concepts

1.1 Principles of Thermal Separation Processes In a chemical production plant, products are produced by the chemical and physical conversion of raw materials or intermediate products. The production unit is a completely integrated technical operating unit on the site. It is connected with other units on the site by transportation and personnel routes, and pipelines for raw materials, auxiliary substances, products, utilities, and energy. It usually consists of the actual production unit and several off-site facilities, as shown in Fig. 1-1. The main unit contains the unitprocesses and operations, such as separation, combination, division, formulation, heat transfer, conveying, storage, packing. Figure 1-2 shows a general set-up which is independent from the type of process. The combination of unit processes and operations with respect to product properties depends on the product produced. During the chemical conversion of raw materials, homogeneous and heterogeneous mixtures (Figs. 1-2 and 1-3) are generated. Both reactants and products may be found in these mixtures, according to the yield and conversion of the chemical reaction. By means of thermal separation processes these mixtures must be treated to obtain the desired products to a demanded purity and to enable the raw materials to be recycled. Processes to separate physically homogeneous (one phase) and heterogeneous (two or multiphase) mixtures are listed in Table 1-1. The driving force of the separa-

tion process usually forms the criteria for the separation. Homogeneous mixtures with a molecularly dispersed distribution of individual components may only be separated by means of a thermal separation process. Thermal separation processes are mass transfer operations, driven by molecular forces. Mass, and often heat, is exchanged between at least two phases of different composition. The phases are the mixture phase(s) and a selective auxiliary phase. The auxiliary phase is generated by either adding heat and/or by means of an auxiliary substance. The required driving forces, concentration, and temperature gradients, are formed due to the auxiliary phase. In Fig. 1-4 thermal separation processes are listed and are denoted by the phases contributing to mass transfer in Table 1-2. Thermal separations of mixtures are carried out in the following individual steps :

Step 1: An additional phase is generated by supplying energy to the system, or by adding an auxiliary component. - Step 2: Mass, and often simultaneously heat, is exchanged between phases. This is achieved by the addition or removal of energy. - Step 3 : After completion of the interchange process, the phases are separated. Together with the separation of the phases a (partial) separation of the mixture occurs. -

All thermal separation processes follow this order of events. The basic principles of thermal separation processes are now formulated and will be discussed in detail.

2

1 Basic Concepts

Main plant Process consisting of physical and chemical unit operations to produce desired products Off-sites, auxiliary equipment

- process control of the main plant

control room sometimes with a process control computer, control devices for drives, production lab, instrument air station

Inputs Energy

-+

-+

- supply of energy to main plant, generation and distribution

--*

Excess energy

- transport to the process unit of the raw material and auxiliary

-+

materials, transport of the products (roads, rail connections, harbor)

Main products

-+

By-products

+

Waste

of electrical power, heating system for heating media such as

hot water, steam, dyphil, salt melts - provision of auxiliary materials (adjuvants) such as heat

Raw materials Auxiliary materials

-

transfer media, coolants, catalysts, solvents, inerts ~~~

~_~

~~

- storage of raw and auxiliary materials, and products, spare parts, tools and materials for repair work and maintenance

_~

-

__ __

- disposal

treatment of waste gas and wastewater, reprocessing of solid residue and waste disposal

+ Waste gas

Wastewater

- facilities for the operating personnel

Fig. 1-1. General production process set-up.

Raw material

I

metering, preheating

i

Main product ready for storage or shipment

~

Path of raw material or product

Fig. 1-2. Basic flow chart for the main part of a production plant.

1.1 Principles of Thermal Separation Processes

3

Y 1 Y 1 A

Y

r

Reaction mixture

A b : : +

Phases in Cocurrent Flow (Principle of Parallel Flow)

Phases in Countercurrent Flow (Principle of Counterflow)

Phases taking place in mass and heat transfer are guided in cocurrent flow through the separation apparatus. The maximum efficiency of this separation apparatus is the same as that for a single theoretical separation stage.

Phases taking place in mass and heat transfer are guided in countercurrent flow through the separation apparatus. In this case it is important to disperse the phases with the aid of internals, thereby achieving intensive mixing of the phases. Thus the

.............................

............................ I

I

1 I

: I

Rectification (counterflow distillation) counterflow crystallimtion from a melt Counterflow sublimation Counterflow liquid-liquid extraction Fractionating, Counterflow processes

i

1

Partial distillation Partial condensation Absorption Extraction Adsorplion Cristallization from a solution Drying

~

~

phase transformation

Thermal \epnratioii

I_._________..-.....

L

I

Procesres with auxiliary materials

Absorption Adsorption Convective drying Carrier distillarion Extractive distillation Extraction Entrainer gas sublimation

Principle of classification. Distillation Partial condensation Crystallization Drying (except convective drying) Vacuum d h m a t i o n

Fig. 1-4. Summary of thermal separation processes.

,_r -- -- -_,l

Simple phase transfomiation t) fractionating

Auxiliary product: required for separation tf not required tor separation

4

1 Basic Concepts

Table 1-1. Summary of separation processes. Classes of separation processes

Driving force of separation process

Nature of mixture

Separation processes

Mechanical separations

Gravity Centrifugal force Pressure

Heterogeneous

Sorting (s - s) Dense-media separation (s - 1) Flotation (s - 1 - g) Sedimentation (s - 1) Filtration (s - 1) Pressing (s - I) Centrifugation (s - 1) Hydrocyclone separation (s - 1) Classification Sieving (s - s) Air classification (s - g) Hydraulic classification (s - 1)

Membrane separation

Pressure Electrical field Concentration gradient

Heterogeneous Homogeneous

Ultrafiltration (s - 1) Reverse osmosis (hyperfiltration) (s - 1) Dialysis (s - 1) Electrodialysis (s - 1) Electrophoresis ( s - 1) Permeation (1 - 1, g - g) Gas diffusion (g - g)

Electrical separation

Electrical field

Heterogeneous

Electro osmosis (s - 1) Electrical dust removal (s - g)

Magnetic separation

Magnetic field

Homogeneous

Magnetic separation (s - s)

Thermal separation

Concentration gradient Temperature gradient

Homogeneous

Distillation (1 - 1) partial condensation (g - g) Absorption (g - g), (A) Adsorption (g - g, s - I), (A) Chromatography (g - g, 1 - 1) Extraction (s - s, 1 - I), (A) Sublimation (g - g) Crystallization ( s - 1, 1 - 1) Drying (s - 1) Thermal diffusion (g - g, 1 - 1)

~

Abbreviations: s solid, 1 liquid, g gas to characterize the state of the components of the mixture to be separated, (A) thermal separation process with auxilliary component.

maximum possible interfacial area (phase boundary) for mass transfer is obtained and, hence, the highest possible mass transfer coefficient values. Figure 1-5 shows a “separation column” with stages connected in series in which the key component i of a

mixture is exchanged from the heavy phase to the light phase. Both phases may contain all components of the mixture. A closer inspection of stage n shows that the heavy phase, with a mole fraction of x,, is in contact with the light phase with a

5

1.1 Principles of Thermal Separation Processes

Table 1-2. Characteristics of thermal separation processes by the phases in which mass and heat transfer occurs. ~~

Phase Phase All components Not all components are in both phases 1 2 are contained Phase 2 One (several) comin both phases Phase 1 pure pure ponent (s) is (are) in both phases ~

Immiscible phases in contact

Miscible phases in contact

g

1

g

S

Distillation Partial condensation

Concentration of solutions

Gas drying

Counter current sublimation

g

g

1

1

Liquid-liquid extraction

1

S

Crystallization from a melt

S

S

g

g

Thermal diffusion

1

1

Thermal diffusion

S

S

~~

Absorption Desorption by stripping Adsorption Drying

Crystallization from a solution

Solid-liquid extraction (leaching) Adsortpion

-

Abbreviations: g gas phase, 1 liquid phase, s solid phase.

mole fraction of Y , , - ~ .If x, and are not phase equilibrium concentrations, the fed phases of stage n are not in phase equilibrium, and mass and heat transfer take place. The key component i becomes enriched in the light phase up to a final concentration y,, while the heavy phase is reduced in component i from x,,to x , - ~ . With stage n as a theoretical separation stage, the leaving phases are in equilibrium and no further mass or heat transfer is possible. Therefore, y, and x , - ~ are phase equilibrium concentrations. The heavy phase, with concentration x,, - arrives at stage n - 1 and comes into contact with the light phase, with concen-

tration y n P 2 .An exchange, similar to that in stage n, takes place. The discussed example shows that for countercurrent phase flow, single stages are connected in series in one separation apparatus. The light phase leaving a stage is guided to the following stage whereas the heavy phase is guided to the previous stage. A theoretical stage is that part of a separation apparatus where mass or heat transfer take place in which entering phases are not in phase equilibrium, while the leaving phases have reached phase equilibrium (see Chapter 1.4). In a practical separation stage, equilibrium is often not achieved. The efficiency

6

1 Basic Concepts

transfer ratio, depending on whether y is only locally valid or constant across the cross section of the column.

Phases in Cross Flow (Cross Flow Principle)

LP

HP

L-l---r Fig. 1-5. Countercurrent flow of two phases in a separation apparatus. n - 1, n Stages connected in series LP Upflowing light phase HP Downflowing heavy phase Molar fraction of the key component in x the heavy phase Molar fraction of the key component in y the light phase

Phases taking part in mass and heat transfer flow across through the separation apparatus at an angle of 90" to each other. The separation efficiency depends on the equilibrium location and the ratio of the phase fluxes, but is often low in an individual separation stage. To separate a mixture and obtain pure products, several separation stages are connected in series. This is done most effectively with countercurrent phase flow. Phase cross flow and parallel feed of one phase to individual separation stages are sometimes used. However, cocurrent flow is of non importance.

compared with a theoretical stage is ex- ~i~~ Requirement pressed as the stage efficiencyfactor, E (exchange ratio, enrichment ratio, MURPHREE The time needed to separate a mixture in a efficiency) (Fig. 1-5): discontinuous operation is the effective residence time. For continuous operation, it is separation effect of a practical stage the mean residence time t,, of the mixture E= in the separation apparatus: separation effect of a theoretical stage E=

Yn-Yn-1 Yn*-Yn-1

(1-1)

t,

V

=T

V

where

where

y,* - y,-,

possible theoretical enrichment of the key component in the light phase ( y ; phase equilibrium concentration at x, y , - y n p 1 actual enrichment of the key component in the light phase

I/

This often has to be distinguished as a local transfer ratio, as opposed to an overall

Short-, medium- and long-term separation processes can be distinguished depending on the time requirement:

filled volume of the mixture in the separation apparatus (determined by the volume of the apparatus and the degree of filling) effective volumetric flow of the mixture

1.2 Thermal Separation Process Modes

Short-term processes (t, < 30 sec). Examples: Spray drying, gas adsorption, precipitation crystallization. - Medium-term processes (30 sec< t, < 2h). Examples: Absorption, rectification, drum drying, pneumatic-conveyor drying, sublimation, extraction, crystallization, liquid adsorption. - Long-term processes (1 h < t , < 1 d). Examples : Rotary drum drying, vacuum tumbling drying, vacuum freeze drying, fractionation crystallization. -

Energy Supply For the thermal separation of a mixture in an apparatus, energy has to be supplied in the form of: -

Heat, to increase the sensible heat of the flowing masses and to supply latent heat.

-

7

Flow energy, for pressure drops in the ap-

paratus and the connecting pipework. Mechanical energy, for example for dispersing, pulsing, stirring and pump circulation devices. - Work, to operate peripheral machines such as compressors and vacuum pumps. -

1.2 Thermal Separation Process Modes Apparatus for the thermal separation of mixtures may be operated both discontinuously (intermittently, batch production, stagewise operation) and continuously (steady-state). In the following section, the operating modes are briefly illustrated. The advantages and disadvantages are listed in Table 1-3.

Table 1-3. Comparison of continuous and discontinuous operation to achieve the same separation

problem.

Comparison criteria

Operating mode Continuous

Mathematical description of the separation process, modeling Investment cost of separation unit Operating cost of separation unit Operation of separation unit Automatic control of separation process Working stress on unit components Environmental pollution, possibility of accident Operation reliability, flexibility in the case of breakdown of separation unit parts, safety buffer Flexibility to adjust to other mixtures to be separated

Discontinuous

Simpler Less Less Easier Possible with less expense Less Less

Higher Better

8

1 Basic Concepts

Continuous Operation : In continuous operation the mixture being separated is continuously fed to the separation device. It is continuously separated into two or more fractions, which are continuously withdrawn from the separation device. An ideal binary mixture can be separated into almost pure components in a separation column operated continuously with countercurrent flow. To separate a mixture of k components, k - 1 columns connected in series are needed. Discontinuous Operation: With discontinuous operation the mixture being separated is charged to the separation device. During a time period, the “batch period”, the mixture is separated mainly into two fractions of defined different compositions. One fraction is continuously withdrawn from the separation device, while the other remains in the device and is withdrawn at the end of the batch time. Discontinuously operated processes mainly in one stage - allow incomplete separation of a mixture; the obtained fractions are treated in subsequent stages (this is the case for multistage discontinuous separation). Alternating Operation : If in a separation apparatus after a loading process (separation of a mixture) an unloading process (the regeneration of a substance aiding separation) is required, at least two sets of equipment are operated alternately. Therefore, steady separation of the mixture is guaranteed. In the case of the adsorption of a substance from a gas phase in a container adsorber (see Chapter 4), for example, a solid

[

f s u m of amount) entering the system (transport)

+

[

adsorbent adsorbes adsorbate (key component in the gas phase). Adsorption continues to an upper loading limit. After the maximum load has been reached in the first adsorber, operation is switched to the second adsorber. The loaded adsorber is then regenerated by dampening, drying and cooling. After the regeneration cycle is finished, the adsorber is ready for loading again.

1.3 Mass Balance, Energy Balance, Exergy Balance In general, the first step in the design of a separation plant is the balancing of individual apparatuses and parts of the plant. Balances are done with respect to energy and mass fluxes, in connection with a schematic representation of the process (flow diagram).

1.3.1 Mass, Energy and Heat Balances The balancing of chemical engineering systems follows the sequence listed in Fig. 1-6. Of the variables listed for process design, mass and energy (usually in the form of heat, enthalpy, and exergy) are of most interest. These variables may also be used for planning, evaluation of systems, analysis, and synthesis. Based on the laws of conservation of mass, energy and momentum, balance equations are set up [1.1] - [1.5]. For a general open system

f sum of amount ) generated in = the system (transformation)

1

f s u m of amount) leaving the system (transport)

1 f

-k

increase of \ mass stored in the system (accumulation)

1.3 Mass Balance, Energy Balance, Exergy Balance

9

Determination of the balance area of the system by defining real or imaginary boundaries (unit complex, unit, part of the unit, individual apparatus, part of an apparatus, volume element)

Determination of the balance size of measured or valuable system properties (total mass, component mass, atomic mass; state variables such as enthalpy, entropy; momentum; cost)

Process description Flow sheets

4l

Formulation of balance eauation to obtain a list of all balance variables of interest in a quantitative form

1

+

~

Solution of the balance equations

I Optimization

Environmental requirements Safety provision requirements Safety requirements

Dertermination of process --* control equipment

+

Precalculation and economical calculations

Fig. 1-6. Balancing of processes (schematic simplification of the concepts) [1.1].

Depending on the problem or task a integral or differential balance equation is generated : 0

Differential, to investigate a process in a differential volume element or at an interfacial surface element

0

Integral to determine the streams entering and leaving the system

Differential balance equations lead to velocity, concentration, and temperature profiles in the system, or at the boundary surfaces, after solving the corresponding differential equation system with suitable bound-

10

1 Basic Concepts

ary conditions. Integral balance equations give a basis for evaluation of the total system with respect to energy and mass. Results of integral balance equations are often presented in a table or in a flowchart (product and energy scheme, mass and heat flowchart, etc.).

The material balance for an individual component k is

c i

-

-

m i j , * Wi,k,a

c mi,, i

*

m , k =

wi,k,w

+ mS,k

(1-4)

where Mass Balance

mi,a

Material balances (mass or quantity balances) can be general or total material balances over the complete system and must be distinguished from material balances for individual mixture components (Fig. 1-7). Using the terms in Fig. 1-7, for an open system the general integral balance equation is

wi, k, a 9 wi, k, o

c mi,a+ i

mQ =

c i

+ m,

(1-3)

In steady-state operation (continuous feed and withdrawal of material without start up and shut down procedures) the accumulation terms mQ and m, are not required.

m i ,w

rfzp ms, I j z Q , k ,

mass flux i (a feed, o product) mass fraction of component k in stream i ms,k intensity of sources (Q) and sinks (S)inside the balance area generally with respect to component k

m Q , k and m $ k take, for example, chemical reactions involving k into consideration. If k is a reactant in j single reactions taking place simultaneously in the balance area, for steady-state operation,

m l , a ~W l . k , a , '1.a

BA

-BB

-

Fig. 1-7. Balance scheme to derive the balance equations. BB Balance boundary BA Balance area

1.3 Mass Balance, Energy Balance, Exergy Balance

11

the balance area and uj,k is the stoichiometric ratio of k in the reaction j . dnk Instead of mass and mass fractions, it is 'j = (1-6) convenient to use another kind of substance Vj,k * I/. dt flow or concentration units to set up mateis the equivalent reaction rate of the reac- rial balance equations. Useful conversion tion j , V is the volume of reaction mass in relationships are given in Table 1.4. where

Table 1-4. Conversion between concentration scales of a mixture component.

Mass fraction wi

Molar fraction

Partial density

xi

ei

Mole ratio of component i ') x m ,1

Mass fraction m;

wi

Molar fraction

ni x . = __

C.i i

c-M; ei

e

i

Partial density

Wi'

e

ei

xi * Mi M -

e

Molar ratio of component i in the inert or carrier phase

I) Reference is the mass of the remaining mixture excluding component i, expressed as carrier mass m 1

c mi

2,

')

i

Density e = - @ = L e i V i n. Molar ratio carrier load X i = n1

Used variables: mi mass of mixture component i number of moles of mixture component i

ni

M =

_f

c xi Mi mixture mean molar mass of the i

9

V

total volume of the mixture

ml

mass of reference component 1

c mi total mass of the mixture c total number of moles of the mixture i

i

ni

12

1 Basic Concepts

For a complete mass balance over a balance area at steady-state, in which only I physical transformation of matter occurs, an equilibrium system consisting of Eq. I (1-3) and k - 1 equations (1-4) for k active components over the balance area, must be where set up. Due to the valid stoichiometric con- hi,,, hi,, dition, for the individual phases, the kth equation (1-4) gives Q cwk=l

(1-7)

k

Ways to formulate material balances for differential volume elements as the balance area, and methods to solve this differential equation system may be found in, for example [1.1] and [1.3]. Energy and Heat Balances

c Ei,a+ EQ i

= E,

+ c Ei,, + Es

(1-8)

i

where E . . E,,E,

Ei,, EpEs

.

QQ,Qs

specific enthalpy of mass flux ki entering and leaving additional heat flow supplied heat flow lost intensity of heat sources (Q) and sinks ( S ) in the balance area

QQ,Qs account for the exothermic and endothermic nature of phase transformations and chemical reactions under steady-state conditions. If, for example, over the balance area, chemical reactions involving component k occur QQ- Qs = - v * h/r,

Based on the law of conservation of energy, an energy balance analogous to the material balance may be set up for any bounded balance area. Using the terms from Fig. 1-7 for the energetic and materialistic open system, the energy balance becomes Ea+

Qu

(1-9)

includes all energy forms, such as potential and kinetic energy, binding energy, heat energy flow through the system boundary, supplied (a), or removed (0) energy supplied or removed by the mass flux mi intensity of the heat sources or sinks inside the balance area

In process design, a heat balance is often sufficient. From the first law of thermodynamics, for the balance area in Fig. 1-7, the mthalpy or heat balance equation is

- c rj - vj, - hhR, = j

=

- u k Ai;,

*

*

Wi,k,a

(1-10)

i

where AhR,j reaction enthalpy for component k in reactionj AhR the total reaction enthalpy for component k (ALR> 0 for endothermic reactions, Ah;, < 0 for exothermic reaction) U, total conversion of k hfk molar mass of k

1.3.2 Exergy Balance When the economics of plants are considered, the cost of raising primary energy causes a domination of energy costs over investment costs. The optimum cost of a process corresponds to minimum use of energy. Therefore, it is necessary to investigate processes which have high energy consumption in order to optimize energy usage.

1.3 Mass Balance, Energy Balance, Exergy Balance

To evaluate the energy utilization, energetic and exergetic process analyses are used. Since exergetidanergetic flowcharts show local internal and external irreversibilities, the locations and quantities of heat losses may be detected, leading to thermodynamic optimization from the consideration of process energy improvements. According to Fig. 1-7 the exergy balance equation, for a steady-state open system, under isobaric conditions, is

=

GE,w+

c

___

dQ,

+

+

G,,w AG, i

(1-11)

where exergy of input Ea and output E, energy fluxes Gj,a,Gi,w exergy of input Yizi,a and output m;,w mass fluxes Qm heat flow supplied G E , a , GE,@

G; = Yiz; * [(hi - h,) - T, . (s; - s,)]

entropy temperature, enthalpy and entropy, referred to surrounding conditions or to a particular system state reference exergy losses due to irreversibilities

AG,

AG,

(1-12)

=

T, * AS,

(1-13)

More information on energetic and exergetic analysis can be found in the literature [1.6- 1.91. The exergetic analysis of a rectification unit is given as an example in [1.10].

13

1.3.3 Calculation of Balance Equations Calculation of mass, energy and exergy balances for the total separation plant are usually done sequentially from apparatus to apparatus. They are occasionally carried out simultaneously by an iterative method, considering the corresponding equation system for the total separation process. In this case, it is necessary to develop a calculation flowchart with coded interface and ramification. The process structure is conveniently presented graphically. After mathematical formulation of the process, the number and values of the independent system variables are determined. Finally, the balance equations are solved sequentially. To calculate the balance data using a computer [1.1, 1.11-1.161 a flexible programming system is required. The main program controls and organizes by assigning priorities to the calculations via references to the corresponding process steps, mass fluxes, and computation parameters. It organizes the intermediate storage of calculated mass fluxes and state variables, and transfers these as input variables to the following stages. Initial variables for each stage are calculated in subroutines, based on stage specific theoretical and empirical models. By including the graphical methods in the balance, economic optimization by, for example, minimizing the energy flow [1.17] or optimization with respect to complete safe operation of the process can be performed [1.18]. It is often convenient to present the results for a balance over a single piece of equipment, a process unit or total processes, and, if necessary, for combined processes. They can be quickly understood in clearly arranged flowcharts (SANKEYdiagram).

14

1 Basic Concepts

1.4 Phase Equilibria 1.4.1 Basic Concepts With thermal separation of mixtures, usually in open systems, an exchange of heat and mass occurs at the phase interface. When phase equilibrium is reached, no further heat or mass transfer takes place. Basic concepts and equations are now introduced [1.20], in order to describe phase equilibria:

- Thermodynamic Systems: Any quantity

-

-

-

-

-

-

of matter that is separated from its surroundings by rigid or imaginary boundaries and whose properties may be unequivocally and completely described by thermodynamic, macroscopic state variables. Open Systems: Thermodynamic systems in which matter and/or energy may transfer with the surroundings, through the boundaries. Closed Systems: Thermodynamic systems in which only energy may be exchanged with the surroundings. The system is closed, containing constant mass, but is not isolated. Adiabatic (thermally) Isolated Systems: Thermodynamic systems in which no heat or mass transfer with the surroundings takes place, although exchange of energy in other forms, e.g., shaft work, is possible. Closed (isolated) Systems: Thermodynamic systems in which no exchange of energy or matter occurs with the surroundings. Homogeneous Systems: Single phase systems with identical macroscopic properties in each volume element. Heterogeneous Systems: Two or multiphase systems with at least one abrupt change in macroscopic properties at the boundar(y/ies) between the phases.

Phase: A physically homogeneous region of a system, contained within a phase boundary. Each volume element in one phase has identical macroscopic properties. The change of state variables, i. e., those not dependent on mass, such as temperature, pressure, composition, etc., are continuous and time independent (without, for example, a step change). - Dispersed (discontinuous) Phase: A phase consisting of homogeneous matter, which is scattered in space, and dispersed in the continuum of the other mainly coherent phase. - Continuous Phase: A uniform, nondispersed phase. - Free (Gibbs free) Enthalpy (thermodynamicpotential), G : The relationship between the enthalpy H, entropy S and internal energy U is defined as

-

G=H-T.S or G = U + p . V-T-S

(1-14)

Therefore, the Gibbs free enthalpy is that part of the enthalpy which, at a reversible change of state, can be converted into other types of energy. For a closed system (constant mass or for a pure phase), the total differential for the free enthalpy d G is dG =

(z)p (g) - dT+

T

.dp

(1-15)

and, from Eq. (1-14) dG=dU+p*dV+ V*dp-

T-dS-S. dT

(1-16)

For a reversible change of state it follows from the first law of thermodynamics that: d U -tp . dV = T . dS

(1-17)

1.4 Phase Equilibria

(dG), = ,ul dn,

and therefore dG= V-dp-S-dT

(1-18)

With constant mass, the free enthalpy is only a function of pressure and temperature. Comparison of the coefficients in Eq. (1-15) and (1-18) gives:

(g)p= -S

and

(5)

=

V (1-19)

T

For an open system, the free enthalpy is not only a function of pressure and temperature, but also a function of the amount of mass n of each individual component. G = G(T,p,n1,n2,...)

(1-20)

The change of free enthalpy is therefore given by

.dnl +

an2 p,T,n,,n3 ... + (E)

By integrating this equation under the assumption of constant pressure and temperature, the free enthalpy of the system shows a definite dependency on the composition and chemical potential of the components of the mixed phases in the system. G = p l . n , + p 2 - n 2 +...

(1-24)

For a change of chemical potential it follows from Eq. (1-24) that (dc),,, = n1 dp,

+ n2 - dp2 + . . .

(1-25)

and with the requirement of dG = 0 at equilibrium,

n, . dp, + n2 - dp,

+ . .. = 0

(1-26)

or

This is the Gibbs-Duhem equation in two different forms and is of fundamental significance in proving the consistency of precalculated or experimentally determined equilibrium data.

. dn2 + . . .

hence dG = - S . d T + V . d p

(1-23)

(1-27)

d G = - S * d T + Vedp-t p. T , n 2 . . .

+ p 2 . dn, + . . .

15

+ (1-21)

For individual components the isothermal and isobaric change of state is

or by introducing the chemical potential p of the individual components, we can write

Chemical Potential (partial molar free enthaipy): An expression for the change of the free enthalpy of a system, which consists of a mixture, if 1 mole of the mixture component i is added to a n infinite amount of the mixture. Hence,

(1-28)

In a mixture the chemical potential of a component i can be expressed as

16

1 Basic Concepts

where poi is the standard chemical potential, i. e., the chemical potential of the component i at the standard state. For any component of a gas mixture, the standard state is the same as for the standard state of a pure gas component at the temperature and pressure of the system. For any component of a liquid mixture, the standard state is the same as the standard state of a pure liquid component at the temperature and pressure of the mixture. The partial molar free mixing enthalpy A& is

A p i = i ? . T . ln(y,.xi)

(1-30)

-

where y i xi is the activity and yi is the activity coefficient of the component i of the mixture. The activity coefficient here is defined as lim yi = 1

(1-31)

X,'l

Equilibrium: A system is at equilibrium if, in fixed surrounding conditions, no change on a macroscopic scale occurs in the system. Therefore, there is no tendency toward an exchange of matter and energy between the phases. Exchanges of matter and energy between phase boundar(y/ies) of the heterogeneous system are reversible. The equilibrium between the phases (the phase equilibrium) is sensitive to changes in the surrounding conditions. Compositions of the phases at the desired phase equilibrium are independent of time, the amount of material and the direction from which equilibrium is reached. Equilibrium Conditions: Equilibrium between the phases of a heterogeneous system is reached if the following conditions are valid: no local pressure differences exist with respect to time and space (mechanical equilibrium) dp = 0 therefore p1 = prl = . . .

(1-32)

and no local temperature difference exists with respect to time and space (thermal equilibrium) d T = 0 therefore

T, = T,, = . . .

(1-33)

These two conditions meet the requirements to describe an isolated system in a state of equilibrium. The system is stable, when the system entropy reaches a maximum, hence dS = 0

(1-34)

The equilibrium condition for a closed system may be derived from the fact that any infinitesimal change of state at equilibrium is reversible. For example, we have dU=dQ+dW=T*dS-p.dV

(1-35)

or

d H = T -d S i - V* dp

(1-36)

The equilibrium conditions for an adiabatic and isochoric change of state are given by, (dU)v,, = 0

(1-37)

for the adiabatic and isobaric change of state,

(CW),,, =0

(1-38)

for the isothermal and isochoric change of state,

(W" ,=0

(1-39)

and for the isothermal and isobaric change of state,

The internal energy U, the enthalpy H , the free internal energy F, and the free enthalpy G , must be at a minimum for a system to be at equilibrium. This is comparable to a me-

1.4 Phase Equilibria

chanical system where the potential energy is a minimum at equilibrium. Combining Eq. (1-21) and (1-28) gives dG= -S.dT+ V-dp+zpi.dni

(1-41)

The equilibrium condition for an isobaric and isothermal change of state for an open system is (dG),,

=

-

p i dni = 0

(1-42)

At equilibrium, mass transfer between the phases of a system is reversible and a material balance requires (1-43)

dn,,, = -dni,,, and so it follows that p r. , = ~ p. I , I I - Pi,III

=

. ..

(1-44)

Therefore, phase equilibrium is reached when the chemical potential of each chemical species in the system is the same in all phases. For the equilibrium of two phases I and 11 using Eq, (1-291,(1-30) and (1-44) it follows that

Rearrangement defines the “distribution coefficient” (%equilibrium constant’? K:, giving

xi,II= KT * xi,,

(1-46)

and the relationship (1-47) with . - fi or,II . APoi = P or,,

(1-48)

17

Knowledge of the distribution coefficient or the equilibrium relationship = f(xiJ as a function of pressure and temperature is essential for the design of thermal separation processes. The relationship is found in three steps: - Step 1: Evaluation of the activity coefficient yi Of each component in each phase at a different temperature, pressure and composition, measured experimentally or determined from an appropriate correlation. - Step 2 : Determination of the differences in the chemical Potential APoi of each component under the same conditions (see Step 1). - Step 3: Evaluation of KT or the equilibrium relationship x ~ ,=~f ,(xi,,).

1.4.1.1 General Differential Equation for the Equilibrium Between TWO Phases

The criterion for the equilibrium of two phases is p i , , = pi,II or dp,,, - dpi,,, = 0. The general differential equation for the equilibrium of two phases I and I1 may be written:

(1-49) This relationship links the state variables of pressure p , temperature T and molar fraction xi of the component i in both phases I and 11. The enthalpy A ~ Z ~ ,and , , ~ , the volume AK,I,II are partial molar phase transfer quantities. The partial molar quantities of a binary mixture may be determined by graphical method if the appropriate quantities are known as a function of the mixture composition (Fig. 1-8).

18

1 Basic Concepts

q P

vl,p&,p

v

X

Example:

-

Partial molar volume of both mixture components Molar fraction of component 1 Mixture volume

-.Ah;,,,

R *T

R - T2

(1-53)

or dP dT

-

The molar volume Vof a real fluid mixture is

__ -

V = x l . I / , , + x , . T -/ 2 , p = x 1 -. 1 / , p +

where

+ (1 - Xl)

(1-50)

G,p

where q,, and 1/2,p are partial molar volumes of components 1 and 2 of the mixture. The mole fraction x1 of component 1 varies between the limits 0 and 1.

-

q,#=v-x,.=

av

(1-51)

ax1

V + (1 - x,) .

aV ~

ax,

Ah1.1, T -A4.11

Ahi,ii = Ah,,g

AI/,Il =

dp/dT

-

VI

(1-54)

the molar evaporation enthalpy of the components at a given p and T the differences in the molar volumes of the saturated vapor and liquid of the system components, at a given p and T the slope of the vaporpressure curve

(1-52)

In Fig. 1-8 the measured volume of the mixture V is plotted against the molar fraction xl, A tangent from point A of the v(x,)curve gives the points D and E. Points B

1.4.1.2 The Gibbs Phase Rule The Gibbs phase rule describes how many state variables of a muitiphase system may vary independently (degrees of freedom of

19

1.4 Phase Equilibria

the system) without disturbing the systems equilibrium :

P+F=K+2

(1-55)

where

F the degrees of freedom of the system at equilibrium (chosen from the state variables pressure, temperature, concentration of each component in each phase); number of variables describing the state of the system which may be varied independently without disturbing the system equilibrium P number of phases of the heterogeneous system (one gas phase, one or more liquid phases depending on the miscibility of the components, one - at mixed crystal forming - or several solid phases according the number of crystal types) K number of components; the minimum number of components of the system forming the phases which must be independently declared

F = 0 invariant system F = 1 monovariant system F = 2 divariant system, etc. The phase rule is important for thermal separation processes as, if certain process parameters are choosen, it establishes which state variables are cogently fixed at an arbitrarily adjusted phase equilibrium (Table 1-5).

1.4.2 Liquid-Liquid Equilibrium 1.4.2.1 The Nernst Distribution Law

For a dissolved substance S in two nonideal, immiscible, liquid solvents such as T and L at constant pressure and temperature the distribution according to Eq. (1-45) is

pas,, + R . T.ha,,,

=

+ . T . lnczs,II

(1-56)

= pos,II R

where as is the activity of S. For phase equilibrium of two liquid phases, the standard state of both phases is the same (normalized to pure liquid S). It follows that POSJ

(1-57)

= POS,II

and therefore including Eq. (1-56) as,1 = as,II

or

YS,I * *S,I = YSJI

. XS,H

(1-58)

To describe the concentration dependance of the activity coefficients Y , , ~and Y , , ~ ~the , computational methods shown in Table 1-8 may be used, for example the NRTL or UNIQUAC methods. For a substance S distributed in phase I (raffinate R) and phase I1 (extract E) with a small concentration of cS,, and c ~ , the ~, Nernst Distribution Law is an approximation given by

A substance is distributed between two liquid phases so that the same concentration ratio appears in both phases. It is independent of the total amount of phases at constant pressure, constant temperature and similar molecular forms in each phase. For separation by liquid-liquid extraction, the Nernst distribution law describes the equilibrium between raffinate and extract phases if the carrier component T and solvent component L are not miscible (see Chapter 6). Table 1-6 shows additional variations of the Nernst distribution law.

5. Selected thermal separation processes. Examples of the Gibb’s phase rule. separation Phases

on o f binary

Number Type 2 1, g

2

quid n

2

orption

2

orption

3

on of le binary , carrier on of one nt

21 Ig

1, g

s, g

1, 1

Components Number Type 2 Miscible liquids

2

3

3

3

Immiscible liquids

Degrees of freedom 2

1

3

Carrier Key component Solvent

3

Inert gas Adsorbate Adsorbent

3

Inert gas Absorbate Solvent

Process parameters Pressure, concentration o f one component, e.g., in the liquid phase Pressure

Consequences for the remai variables describing the syst explanations Concentration of the second nent in the liquid phase, the tration of both components vapor phase, and the temper all fixed (boiling point diagr Concentrations and tempera strongly fixed, no simple se possible (boiling point diagr

If the concentration of the ponent in the raffinate rema concentration of the extract fixed (distribution equilibriu

Pressure, temperature, concentration of the key component in the liquid phase

Fixed loading of the compo adsorbed in the absorbent ( tion isotherm)

Pressure, temperature, partial pressure o f the absorbate in the gas phase

Concentration of absorbate liquid phase is fixed (absorp isotherm)

Pressure, temperature, partial pressure of the absorbate in the gas phase

. (continued) separation Phases Number Type

Components Number Type

Degrees of freedom

Process parameters

Consequences for the remain variables describing the syste explanations

Components without mixed crystals load

ion

Dissolved substance Solvent

ration and m llization

Concentration of the dissolv substance fixed by the solubi curve

Pressure temperature

Dissolved substance Solvent

zation from n ooling llization

Solid phase concentration an temperature are fixed (solidu liquidus lines)

Pressure, concentration of one component in the liquid phase

Components forming a mixed crystal

ation from

Moisture content of the prod Pressure, temfixed (sorption isotherm, 1st perature, partial pressure of the period) moisture in the gas

Dry product Moisture Dry gas

pic product ng period)

Dry product Moisture Dry gas

on drying groscopic by hot gas g period)

Pressure temperature

Pressure

Temperature

Fixed saturation loading of t product with moisture (sorpt isotherm and 2nd drying per

Boiling point temperature of saturated solution is strongly by the pressure (vapor-press curve of the saturated soluti Temperature determines the pressure of the sublimable c ponents, corresponding to th sublimation pressure curve

tions: s solid phase (solid), 1 liquid phase (liquid), g gas phase (gas).

22

1 Basic Concepts

Table 1-6. Additional formulae of the Nernst distribution law. Correlations and conversion relationships for the distribution coefficient K . 0

Mole fractions x, y are used as the concentration scales for the distributed substances S Y x

-=

0

K?(tP)

Mass fractions w substance S wS. E ~

wS, R 0

= K:(B)

~ and, w~ ~ are, used ~ as the concentration scales for the distributed

(1-61)

.

Molar ratios X , Y are used as the concentration scales for the distributed substance S Y X

- = K$(B) 0

(1-60)

.

(1-62)

.

Relationship between the distribution coefficient

(1-63) 0

Correlation for the distribution coefficient K: [1.21 K$

A , + A z . wS,R

(1-64)

+ A , . w ~ , +R A,. w;,R

Specific substance constants for Eq. (1-64)at tP = 25 "C. System

I

I1 I11

A,

A2

7.1957 0 1.1888 0 1.5569 0

A3

A4

-1022281.0 50722.0 -3.3775 2.0970 - 3.6840 -0.7462

System examples : I Water/toluene/aniline [1.21] 11 Water/benzol/dioxane 111 Watedmethyl isobutyl ketone (4-methyl-2-pentanone)/acetone

Nomenclature used in Eqs. (1-59) to (1-64):

VE

volume of the extract phase nS,R

CS,E = __ nS'E

vE

'S,E

molar concentration of the substance S in the extract phase number of moles of S in the extract phase

C S , R = __ 'R

molar concentration of S in the raffinate phase

IIS,R

number of moles of S in the raffinate phase

VR

volume of the raffinate phase

1.4 Phase Equilibria

23

mole fraction of the substance S i n the raffinate phase mole fraction of the substance S in the extract phase number of moles of extract or raffinate phase weight fraction of S in the extract phase weight fraction of S in the raffinate phase mass of S in extract or raffinate phase mass of extract or raffinate phase mole ratio of S in the raffinate phase number of moles of carrier mole ratio S in the extract phase number of moles of the solvent molar mass of extract and raffinate phase density of extract or raffinate phase A schematic distribution diagram for liquid-liquid extraction is shown in Fig. 1-9. The Y, X distribution diagram is essential for the design of separation apparatus.

1.4.2.2 Representation of Liquid-Liquid Phase Equilibrium

Industrial extraction process systems with three or four liquid components are common. In addition to a vapor phase two or

X-

Fig. 1-9. Schematic presentation of the liquidliquid phase equilibrium, the loading diagram. EC I Equilibrium curve with a constant distribution coefficient EC 11, Equilibrium curve for a concentrationEC I11 dependent distribution coefficient P Plait point Y = X , 45" line AL Mole ratio of S in the extract phase Y Mole ratio of S in the raffinate phase X

three pairs of liquid phases are possible. Now equilibrium of liquid-liquid phases for ternary systems will be discussed in more detail for cases where partial miscibility of carrier liquid and solvent cannot be neglected, and presentation by means of an equilibrium diagram similar to Fig. 1-9 is insufficient. A practical way at constant temperature and pressure to graphically describe the data of phase equilibrium of a ternary system uses an equilateral triangle (Gibbs' triangle). Figure 1-10 shows an equilateral triangle representing a three component mixture. Each apex of the triangle corresponds to one of the pure components T L and S. Any point on the sides TS, TL and LS char-

24

1 Basic Concepts XTXL-

0 0

(1-65)

XS‘l

S

In the case of a liquid-liquid extraction operation, the solvent L must be chosen such that it is very miscible with the solute of the original binary mixture T/S yet immiscible with the carrier liquid 7: Simple extraction processes for a common type of ternary mixture with components T, L, and S are presented in Fig. 1-11. A vapor phase and two liquid phases (in the two phase region) coexist in the heterogeneous system. S is completely soluble in T and L. In the area between A and B, T and L are insoluble in each other. The binary mixture of A and B, x1.l XT= 0 according to point M, will separate into two XL: 0 XLXL- 1 liquid phases of compositions A and B corxs.0 XS‘ 0 XT responding to the mass ratio m/MB. Until for example, the substance S is Fig. 1-10. Equilateral triangle (Gibb’s triangle) for the representation of ternary mixtures. added isothermally to the binary mixture, A Binary mixture state point with molar according to point M, the ratio m M / B fractions of x, = 0.7 and xL = 0.3 will change if the liquid phase with the B,L State points of two ternary mixtures highest concentration of S disappears. Now M State point of a ternary mixture the ternary system is homogeneous. This is resulting from the mixing of B and C xL,xT,xsMolar fractions of L, T, and S in the true at point M’, the saturation point. Similarly, this is valid for points P, P’ and Q, Q‘. ternary mixture at state point M Curve A, M , P’, Q’, B is the line connecting the saturation points or the solubility curve, solubility isotherm or binodal curve acterizes a binary mixture. Point A, for ex- (Fig. 1-11). ample, is a two component mixture with a Below the binodal curve the ternary syscomposition of 0.7 parts of T and 0.3 parts tem is heterogeneous; elsewhere it is a hoof L. Any point within the triangle repre- mogeneous solution of one liquid phase. A sents the composition of a ternary mixture, mixture, such as R in Fig. 1-11 b, below the for example B, C and M; the graphical binodal curve will form two conjugate liqmethod to derive the fractions xT, x, and uid phases, represented by points C and D x, in Fig. 1-10 will now be explained in on the binodal curve. The connecting line more detail. between C and D is the tie line and deIf n, kmol of the ternary mixture at scribes the state of the phase equilibrium of point B is added to n2 kmol of the ternary the conjugate liquid phases. The more submixture at point C, the composition of the stance S the conjugate liquid phases connew mixture lies on the line This is tain, the shorter the tie lines. The equilibshown by means of a material balance. The rium points C and D move toward each location of the mixing point M follows the other on the binodal curve to finally reach lever, or mixture, rule such that the “critical point” P, the plait point.

-

m.

1.4 Phase Equilibria

A 9 z const

a)

b)

25

S=const

A

o'/ I

TA

i

Q

I

i

I

P

\

\

M

BL

Definition of the binodal curve [binode) A Q' P' M' B Binodal curve

Fig. 1-11. Representation of a liquid-liquidphase equilibrium in an equilateral triangle. a) Definition of the binodal curve (binode)

AQ'P'M'B Binodal curve b) Interpolation procedure to determine the tie lines CPD ... Binodal curve CD Tie line PK, K,K3 Conjugation line

If, for example, several tie lines are found experimentally, additional tie lines may be constructed via the interpolation procedure, linking conjugate liquid phases in phase equilibrium, shown in Fig. 1-11b (for additional interpolation methods, see [6.1]). The conjugation line connects the intersecting points that are generated by the lines parallel to the triangle axes TS or E. The starting point occurs at the intersection of the tie lines and the binodal curve. The value of K* for horizontal lines is unity, corresponding to the Nernst Distribution Law and according to Eqs. (1-59)-(1-63). For tie lines inclined to the horizontal the coefficient values are K* > 1 or K* < 1. It is uncommon in extraction processes to find ternary systems of T, L and S where

I Interpolation procedure t o determine the tie lines C P D . . . Binodal curve CD Tie line P K, K, K, Conjugation line

two partially immiscible substance pairs exhibit a miscibility gap at equilibrium as shown in Fig. 1-12. To simplify the procedure of solving extraction problems, equilibrium curves may be transposed from triangular coordinates to rectangular coordinates [0.1, vol. 2, p. 5461. Figure 1-13 shows different graphical methods used to represent the distribution equilibrium. An increase in temperature usually leads to an increase in mutual solubility. The area of heterogeneity under the binodal curve decreases and the tie lines may also change their slopes. This is shown schematically for a liquid-liquid phase equilibrium in Fig. 1-14. The influence of pressure is negligible with regard to technical accuracy.

26

1 Basic Concepts

Closed system with miscibility gap between L and S

Open system with miscibility gap between L and S

Open system with miscibility gap between T and S

System with decomposition of mixture into three phases in region I11

Fig. 1-12. Phase equilibrium of ternary systems with two partially inmiscible substance pairs. I One phase region

I1 Two phase region

111 Three phase region Type A Closed system with miscibility gap between L and S Type B Open system with miscibility gap between L and S Type C Open system with miscibility gap between T and S Type D System with mixture decomposition into three phases in region I11

The liquid-liquid equilibrium data are determined by a straight forward experimental procedure [0.1, vol. 2, p. 5561 or [6.2]. A heterogeneous ternary system is used in the analysis to obtain the tie lines. After intensive mixing, the system is separated isothermally into two conjugate

phases. Both phases are weighed and analyzed and the state points entered on the Cibbs phase diagram. The lever rule for the phases (Eq. 1-65) can be used as a control method. The mixture point R, and the state points of the two conjugate phases C and D form a straight line; the tie line (Fig. 1-11).

1.4 Phase Equilibria al

I

S

27

I

L

X--b

t Y

I

L

Fig. 1-13. Phase equilibrium for a solvent extraction process. a) Closed system of two partially miscible liquids T and L b) Open system of three liquids, T/L and L/S partially miscible R ,,R,, R3 Raffinate phase E l , E2,E3 Extract phase K , , K 2 ,K3 Tie lines BC Binodal curve EC Equilibrium curve P Plait point at K* = 1 Mole fraction of the key component in the raffinate and extract phases x, Y Mole ratio of the key component in the carrier and the solvent X, Y 45" degree line used to transfer equilibrium concentration points x =y

The ratio of the distances CR/m must correspond to the ratio of the mass fractions of the conjugate phases. To determine the binodal curve by the tilration method, the third component must be added slowly to either one of the homogeneous binary mixture, T/S or L / S . If the binodal curve is reached, the mixture becomes cloudy. The analysis gives one point on the binodal curve (point E in Fig. 1-11).

Binodal curves found experimentally may be described by empirical correlations (parabolic approach for system of type C and Hlavaty approach for system of type D in Fig. 1-12 [1.21, 1.221). In the case of a low concentration of S distributed in the raffinate and extract phase, Eq. (1-64) can be used to calculate the distribution coefficient. The activity coefficient in Eq. (1-58) should be calculated according to Table 1-8.

28

1 Basic Concepts 5

S

1

P

L

Fig. 1-14. Temperature effect on ternary liquid equilibria [0.8]. P Plait point T, L, S Components

Data for liquid-liquid equilibria can be found in [1.23-1.301.

1.4.3 Vapor-Liquid Equilibrium 1.4.3.1 One Component Systems The phases of a one component system are usually presented on a p,19diagram. Figure 1-15 shows a p , I9 diagram for water. The diagram is divided into three areas, the solid phase (ice), the liquid phase (water), and the vapor phase (steam). At a higher pressure, due to the existence of additional solid phases, further regions are added (polymorphic). Inside each area the system is divariant; pressure and temperature may be varied independently. The lines separating the regions are the melting or fusion curve, the vaporization

curve and the sublimation curve. These lines generally connect points between the phases at which each of the two coexisting phases are in equilibrium. According to the Gibbs phase rule, a one component system with two phases is monovariant at equilibrium. If the pressure p is changed, then, according to the saturation pressure, the temperature B must also change. The pressure p and temperature 19 are related by the corresponding equilibrium curve. The differential Eq. (1-49) gives the relationship between the two phases at equilibrium. A brief review of the equilibrium conditions, according to Eq. (1-40) gives the following: dgsystem at equilibrium = dgsystem at equilibrium = dgI - dgII =

(1-66)

29

1.4 Phase Equilibria

t I

P 218 Ibarl

From the latent enthalpy of phase transition from I + I I follows

IMC

~

Ah

A S = __

0.987

0.00611

(1-70)

T

Substituting Eq. (1-69), it follows that dp - A h d T T. A V

~

(1-71)

u 0.0075 100

31oc1

37L

Fig. 1-15. p , &Diagram of water. TP Triple point (coexistence of three phases, invariant system) MC Melting or fusion curve (coexistence of solid phase and liquid phase at equilibrium, monovariant system) VC Vaporization curve (coexistence of liquid phase and gas phase (steam) at equilibrium, monovariant system) SC Sublimation curve (coexistence of solid phase and gas phase at equilibrium, monovariant system) C P Critical point p Pressure B Temperature

where dg,, dg,, are the change of free enthalpy in phases I and 11. Substituting into Eq. (1-18), dg1= - S , . d T + q . d p = d g I I = - -311

*

d T + 61. dp

(1-68)

The slope dp/dT of the equilibrium curve between two coexisting phases is therefore

(1-69) where A i is the difference in the molar entropies, and A the difference in molar volumes of phases I and 11.

v

This is the common form of the ClausiusClapeyron equation, also known as the Clapeyron equation. The equation is generally valid and describes all forms of two phase equilibria for one component systems. In the following text, vapor-liquid equilibrium will be discussed. The equilibrium curve is the vaporization line beginning at the triple point and ending at the critical point where the liquid and gas phase are “identical”. Substituting the molar vaporization enthalpy Ah,,g, at the equilibrium temperature T, and the difference of the molar volumes A V of vapor (g) and liquid (I) (TI)into Eq. (1-71), the slope of the vaporization curve is

(6,)

dp

d~

G

-

T.

g

(q,- q)

(1-72)

At a certain distance from the critical point, the molar volume becomes negligible in comparison with GI. Furthermore, if the vapor phase behaves like an ideal gas, may be substituted by RT/p. Eq. (1-72) therefore becomes

q

6,

dlnp ~

dT

Ahl,g %-

R - T2

(1-73)

Upon integration, the following approximation for the vapor-pressure curve is then found :

30

1 Basic Concepts

l n p = - _'his' .R

T

+ const

(1-74)

If Q GI, a plot of Eq. (1-74) as l n p against l/T yields a straight line. This implies the assumption of ideal behavior for the vapor and a vaporization enthalpy independent of temperature and pressure. A better correlation for vapor pressure is as presented by ANTOINE

cal temperature of the components in the mixture. In absorption processes, the temperature is usually higher than the critical temperature of the dissolved gas components. Nevertheless, distillation and absorption processes may be discussed together with respect to relevant phase equilibria. With the equilibrium condition for two phases expressed by Eq. (1-45) or derived from Eq. (1-49), the general equation for gas-liquid phase equilibrium is

(1-75) A, B, and Care substance related constants. Additional methods for computing vapor pressure are listed in [1.58]. If two points on the vapor pressure curve are known, such a s p , , and p z , T,, the vaporization enthalpy may be estimated using the integral form of Eq. (1-73)

1

*

(&

-

":T>

. dp]

P

(1-77)

where mole fraction of component i in the vapor phase pressure of the system at equilibrium p T temperature of the system at equilibrium I? the universal gas constant q,G partial molar volume of component i at pressure p and temperature T yi activity coefficient of i in the liquid phase (see Eq. (1-31)) xi mole fraction of component i in the liquid phase Q L molar volume of liquid component i at temperature T and saturated vapor pressure poi f , fugacity of the pure substance i at temperature T and the corresponding saturated vapor pressure poi yi

(1-76) Methods for computing the vaporization enthalpies are documented and critically examined in [8.1]. Figure 1.16 shows two common presentations of vapor pressure curves for a few different substances. For practical use, vapor pressure data may be found in [8.3, 8.41. 1.4.3.2 Two and Multicomponent Systems The Basic Equation for Vapor-Liquid Equilibrium

The basic equation for the vapor-liquid phase equilibrium forms the major design criteria for apparatus that separate liquid mixtures by distillation or selective absorption. In distillation, the process or actual working temperature is lower than the criti-

The fugacity f , is defined by the chemical potential of a real gas, as postulated by LEWIS:

pi =poi + R . T . lnf,

(1-78)

1.4 Phase Equilibria 1c

[bad

a a’

tP 6

10

[bar] 7 6

5 L

b’

t:

log P

1.c

: 0

0.E

0.7 0.E

0.E 0.4

0.; Fig. 1-16. Vapor-pressure curves of some substances. a) p, T-diagram b) logp, l/T-diagram M Methanol W Water T Toluene E Ethanol B Benzol

31

32

1 Basic Concepts

Fugacity is used to replace the pressure of

an ideal gas pi by a “corrected pressure” A of a real gas i, to produce a universally valid relationship. At zero pressure a real gas behaves like an ideal gas, and fi lim =1 PI’O

(1-79)

Pi

A discussion and derivation of Eq. (1-77) can be found in [0.12, 1.581. Eq. (1-77) is universally valid, and should be applied if real behavior of vapor and liquid phases is to be considered. Considering the influence of pressure on the vapor and liquid properties of a substance, the real saturated vapor pressure p$ is

(1-80)

The molar volume vmay be computed with an appropriate equation of state for v(p,T). Eq. (1-82) can then be integrated. Useful correlations which may be used to calculate state variables are listed in Table 1-7. These equations may be applied to both pure gases and gas mixtures. Most correlations used to calculate the activity coefficients are based on the dependence of the free mixing enthalpy on the activity coefficient for each individual component in the mixture. If two pure components are mixed, the entropy AS, increases. A positive and/or negative change of enthalpy (“heat of mixing”) AH,, and a decrease of the free enthalpy, the free mixing enthalpy AG,, occur. It can be shown that AGM+ T . AS,

AH,=

(1-83)

The free mixing enthalpy Ag,, one kmol of the mixture, is P

Ag, = R . T .

Eq. (1-77) becomes y..p = y;.

xi. p*. 01

(1-81)

AgM = R . T .

For the design of distillation and absorption apparatus, it is necessary to know the equilibrium relationship yi (xi) either as an analytical, homogenous equation, or at least as an equilibrium curve. According to Eq. (1-77), the fugacity or the fugacity coefficients can be calculated or found experimentally to be a function of pressure, temperature and phase composition. An expression for the fugacity is V - E ) . d p

T-

C xi. h a i i

1xi i

*

-

ln(yi x i )

(1-84)

Therefore

Calculation of Vapor-Liquid Equilibrium Data

In- =

=

related to

(1-82)

1xi. lnx, + i

+ E . T.Cxi.lnyi

(1-85)

i

In Eq. (1-85) the first term is the change of enthalpy during the mixing of ideal liquids, a free mixing enthalpy. The second term, change of excess enthalpy AgME, takes into consideration the behavior of real liquids in the mixing process. AgME=R.TCxj*lnyj i

(1-86)

1.4 Phase Equilibria

33

Table 1-7. Method to calculate p, V, T data for real gases and gas mixtures*. Method

Characteristics

Literature

Benedict, Webb, Rubin method ,,BWR equation"

Equation with 8 constants dependent on mixture concentration. Originally developed for light hydrocarbons, but modified and validity region expanded

BENEDICT,M., WEBB,G. B., RUBIN,L. C., .lChem. Phys. 8 (1940) 334. BENDER,E., Habilitationsschrift, Ruhruniversitat Bochum. ORYE,R.V., Znd. Eng. Chem. Des. Dev. 8 (1969) 4, 579.

Redlich-Kwong method

Equation with 2 constants determined by the critical data of a substance, many times extended and improved

Virial equation Wohl method

Correlation for second virial coefficients

WOHL,K., Z. Phys. Chem. 2 (1929) 77

Prausnitz, Gunn method

Correlation for second virial coefficients

PRAUSNITZ, J.M., and GUNN,R.D., AZChE. J. 4(1958) 430.

*

REDLICH,O., and KWONG,J., Chem. Rev.

44 (1949) 223.

Additional methods in [1.70].

Applying a partial mole fraction derivative to Eq. (1-86), the activity coefficients of the components are given by

For example, the activity coefficients of a binary mixture y1 and y2 are

(1-88) and

Some examples for practical use are given in Table 1-8. Each method has been selected and tested on many substances, and their possible applications are given. All of the listed computation methods need not only be used for variables of pure substances, but also for state variables of the mixture. One or more points on the equilibrium curve of a binary system should be found experimentally.

Experimental Investigation of Vapor-Liquid Equilibrium

The procedure to find vapor-liquid equilibrium data by experimentation is as follows: A sample of a liquid mixture, with a known mole ratio in equilibrium with a vaThe methods used to compute A g M L . por phase, is selected and analyzed. Equiincluding Eqs. (1-87)-(1-89), lead to the librium data at constant temperature or methods of VAN LAAR,MARGULES, WOHL, constant pressure can be found. This comREDLICH-KISTER, WILSON, and PRAUSNITPmon procedure is explained in detail by to compute the activity coefficient. HALA,PICKet al. [1.58, 1.591.

34

1 Basic Concepts

Table 1-8. Methods to Calculate Vapor-Liquid Equilibrium Data [1.70, 1.711. Method

Characteristics

Literature

Method of WOHL (VANLAAR,MARGULES,SCATCHARDHAMER)

Effective Volumetric Ratio Polynomial Equation for the excess Gibb's energy, interaction parameters, relatively simple to use

WOHL,K., Trans. Am. Chem. Eng. 42 (1946) 215 [1.13].

Extended Van Laarmethod

For nonpolar binary systems with large dilution in the region up to the critical region

MUIRBROOK, N. K., Dissertation, Univ. of California, Berkeley (1964) CHUEH,P. L., and PRAUSNITZ, J. M., Computer Calculations for HighPressure Vapor Liquid Equilibria. Prentice Hall Inc., 1968.

Scatchard-Hildebrand Solubility parameters and partial Equation, Chaomolar volumes are referred to Seader method 25 "C, suitable for hydrocarbon systems

CHAO,K.C., and SEADER,J.D., AIChE. J 7 (1961) 4, 598.

Wilson equation

Interaction method for the excess Gibb's energy; suitable for totally miscible systems, not applicable for systems with limited miscibility since only the binary parameters are used, applicable to multicomponent systems; only valid for small and medium operating pressures

WILSON,G.M., J. Am. Chem. SOC. 86 (1964) 127.

NRTL equation (non random two liquids)

Method based on the Wilson equation with nonrandomness parameter which can be applied to systems with limited miscibility and nonideal systems; use of binary parameters to calculate multicomponent data; only valid for small and medium operating pressures

J. M., RENON,H., and PRAUSNITZ, AZChE. J 14 (1968) 135. RENON,H., and PRAUSNITZ, J. M., Ind. Eng. Chem. Dev. 8 (1969) 3, 413.

Uniquac equation (universal quasi chemical)

Method based on the principle of the local compositions, similar to the Wilson and NRTL equations, which are derived as special cases; also describes ,,real" liquid phases; only valid for small and medium operating pressures

GMEHLING, J., ANDERSON, T.F., and PRAUSNITZ, J.M., Ind. Eng. Chem. Fund. 17 (1978) 269.

Unifac equation (uniquac functional group activity coefficient)

Contributions of individual functional groups; prediction of the interaction between the activity coefficients of functional groups by interaction parameters, only valid for small and medium operating pressures

J., FREDENSLUND, A., GMEHLING, and RASMUSSEN, P., Vapor-Liquid Equilibria using Unifac. Elsevier Publ. 1977.

1.4 Phase Equilibria

35

n

Fig. 1-17. LABODEST apparatus for the determination of vapor-liquid equilibria. Vacuum, ambient and over pressure operation at temperatures up to 250°C. Presentation according to data of Fischer Labor- und Verfahrenstechnik, Meckenheim near Bonn [1.61]. 1 Flow evaporator with electrical immersion heater Phase mixing chamber to adjust the equilibrium 2 3 Phase contact path 4 Phase separation chamber Solenoid valves to take samples 5 6 Sample take-off, vapor phase 7 Sample take-off, liquid phase 8, 9 Sample take-off, circulation streams 10 Gaseous sampling of vapor phase, i. e., for systems with a miscibility gap

36

1 Basic Concepts

Figure 1-17 shows an experimental apparatus introduced by FISCHER[1.61] to find vapor liquid equilibrium data.

(1-92) In Fig. 1-18, the method used to check consistency is applied to a binary mixture of acetone and trichloromethane. The equilib-

Consistency of Equilibrium Data To check for accuracy, thermodynamic consistency, and lack of contradiction in the equilibrium data found, many methods and criteria using both experimentation and correlations are available t1.62- 1.661. A simple method used to check the consistency of data for a binary system is discussed below: At a constant temperature, it follows from Eqs. (1-21)-(1-26) that (1-90)

a1

I

0

1

10

0.2

x1 0,L

05

0.8

1,O

- 0.1

-0.2

h

- - 0,3

This leads to a simplified Gibbs-Duhem Equation particularly for a binary mixture

The logarithm of the activity coefficients in Eq. (1-91) may be plotted against the mole fraction of component 1. For a consistency check of the values for logy,(x,) and log y 2 (x2)data, the following requirements should be satisfied: Over the complete range of mole fractions the curves should have opposite slopes At the point x1= 0.5 the slope of the curves should be equal but opposite If the curves do not show a minimum or maximum, all y-values should be either larger or smaller than 1 (all data points for both logy(x,) curves should be on the same side of the log y = 0 line) The integral, or rather the area below the log ( y 1 / y 2 )(xI) curve has to be equal to 0

-0.3I Fig. 1-18. Proof of consistency of vapor-liquidequilibrium data for an acetone-trichloromethane system at 1.013 bar. a) log y,,x,-diagram to prove Eq. (1-91) b) log ( y 1 / y 2 ) , x,-diagram to prove Eq. (1-92), 1 Acetone, 2 Trichloromethane

1.4 Phase Equilibria

rium data is supplied by the DECHEMA service for substance properties. The calculation was done with the aid of the Uhde-substance property compiler using the Van-Laar Method. In process engineering the proof of published vapor-liquid equilibrium is of particular interest. In [1.67-1.701 and [1.72-1.851 data collections and bibliographies are listed. Particular attention is given to a collection of vapor-liquid equilibrium data by GMEHLING et a1 and published by DECHEMA [1.72]. Ideal and Real Mixtures

Depending on miscibility and effects resulting from mixing (volume contraction, volume dilation, heat of mixing) of different kinds of liquids, real and ideal mixing behavior and real and ideal mixtures can be distinguished. In an ideal mixture, components are miscible in any mole ratio. The attractive force between different types of molecule, is the same as that between similar types of molecule. During mixing there are no volume effects or heats of mixing. In practice, only approximately ideal behavior is found, for example, in the mixing of substances such as isotopes, optical antipodes, stereo isomers, structural isomers and neighbors in a homologous series. For an ideal mixture, Raoult’s law follows from Eq. (1-81)

y I . p = p . = p O,I. . X i

(1-93)

Definition of Raoult’s law: The partial pressure pi of component i in the vapor phase is at an adjusted equilibrium proportional to the mole fraction xi in the liquid phase. The saturated vapor pressure po,i depends on the equilibrium temperature. For a real mixture the attractive force between different types of molecule is not the same as that between similar types of mole-

37

cule. Molecules are described as similar if they have similar size, structure, and chemical nature. If in a particular mixture, the attractive force between different types of molecule is smaller than that between similar types of molecule, the molecules in the mixture are held together with a smaller force than those in a pure liquid. Thus, the effect during mixing is an endothermic volume expansion and an increase in vapor pressure. There is a positive deviation from Raoult’s law which leads to a pronounced maximum for the vapor pressure, and therefore to a minimum boiling point (for example, ethanol-water mixture). If the attractive force between different types of molecule is larger than that between similar types of molecule, the molecules in the mixture are held together with a stronger force than those in a pure liquid. This results in an exothermic volume contraction during mixing. There is a negative deviation from Raoult’s law, leading to the development of a minimum in the vapor pressure curve, and a maximum boiling point (for example, nitric acid - water mixture). In Table 1-9, different approaches to describe the liquid-vapor (or gas) phase equilibrium are listed. In distillation processes, the operating pressure is usually in the range ca. 0.05-5 bar. It is sufficient to treat the liquid phase as a real liquid, while the vapor phase may be assumed to be ideal. Gas absorption processes are often carried out under relatively high pressure conditions. The real behavior of the vapor phase must be taken into consideration. Representation of the Vapor-Liquid Equilibrium for Binary Mixtures

For the graphical representation of vapor liquid equilibrium of binary mixtures, there

38

1 Basic Concepts

Table 1-9. Gas - liquid equilibrium relationship. Liquid phase Ideal Partial pressure p i of component i in the vapor phase

-

Ideal

PI = Po,1 . XI

Real

1)

-

-

Gas phase

Ideal

PI = Po,1.

Y1

. XI

Real - Real

PI =Po*,[‘ Y1 . XI

Distribution coefficient KT = y ;/x; Relative volatility (separation factor) ai,k Partial pressure p i of the super critical component in the gas phase ’)

Raoult’s Law.

’)

ai,k =

Po. i

Y1 . P0,l Yk’P0.k p I. = H I. . y .l . x l.

ai,k =

Po, k

~

p I. = H I. . x2) 1

~

ai,k =

P,*,i

Yi ~

Yk

*

Po*, k

p I. = p *0,1. . y I. . x I.

Henry’s Law.

are three diagrams of practical importance: the pressure diagram, the boiling diagram, and the equilibrium diagram.

Pressure Diagram: The partial pressures p 1 and p 2 of both components of the mixture, in the vapor phase, and the pressure p, the system or total pressure of the liquid and vapor phase, are plotted against the mole fraction of the liquid phase at constant temperature. Boiling Diagram: The boiling line and dew line are plotted at constant pressure. The boiling line L9(x) connects the bubble points at boiling temperature as a function of the composition of the liquid mixture. The dew line L9(y) shows the condensation temperature Of the saturated vapor mixture as a function of its composition in the vapor. Equilibrium Diagram: For the case of constant operating pressure, the equilibrium composition of the vapor phase is plotted as a function of the liquid phase composition.

Figure 1-19 shows the three diagrams for an ideal binary mixture (for example, a nearly ideal binary system is benzol-toluene). The pressure diagram for an ideal mixture may be described by Raoult’s law. The partial pressures p1(x,), p2(x2)and the overall pressure p(xl) are straight lines. Boiling lines and dew lines for an ideal mixture at a given pressurep may be derived from Dalton’s and Raoult’s laws to give P=CP;

(1-94)

i

and (1-95)

p,I = y 1, .

Therefore, the equation describing the boiling line is

x1(L9,p) =

P - Po,z(d) Po, 1 (8) - Po,2 (19)

(1-96)

and the equation describing the dew line is

1.4 Phase Equilibria

0

x,l%l b)

f

3

100 OC

90

"0

20

XI,

LO

~1160

100

f

-

80 x 1%1

cl

x2

100

8o

Y

[%I 60 Yl

40

20

0'

20

XiLO

60

80

x[%l

-

100

39

Fig. 1-19. Pressure diagram, boiling diagram and equilibrium diagram for the almost ideal mixture benzol-toluene. a) Pressure diagram for a mixture of boiling point temperature 100 "C p 1(xl),p2(x2)Partial pressure lines of benzol and toluene p(xl,xz) Total pressure line X1 Molar fraction of benzol in the liquid phase Molar fraction of toluene in x2 the liquid phase b) Boiling diagram at a total pressure of 1.013 bar I Liquid phase region I1 Steam or vapor phase region 111 Two phase (wet steam) region B Bubble point line D Dew point line L9 Temperature c) Equilibrium diagram at a total pressure of 1.013 bar y Molar fraction of benzol in the vapor x Molar fraction of benzol in the liquid phase -- Conceptional lines to generate the equilibrium diagram from the boiling diagram

Similarly, the equilibrium curve of an ideal mixture may also be derived from Dalton's and Raoult's laws to give Po, 1

yI = P0,l

XI

f

*

XI

(1-98)

P0,2 x 2

The relative volatility a1,2,independent on concentration (see Chapter lS), is (1-99)

Boiling lines and dew lines may be calculated stepwise for a chosen temperature 8. po,l and po,z are the saturated pressures of pure components 1 and 2 corresponding to temperature 8.

y, =

a1,2 * X l

1 + x1 *

@1,2

- 1)

(1-100)

The equilibrium curve may also be developed graphically from the boiling diagram,

40

1 Basic Concepts Fig. 1-20. A schematic of the equilibrium behavior and important thermodynamic functions of a binary system according to Rock [1.861, a) Pressure diagram b) Boiling diagram c) Equilibrium diagram d) Activity coefficient diagram e) Relative volatility diagram f ) Free excess enthalpy of the mixture g) Mixing enthalpy and excess entropy 1st Column: Ideal mixture 2nd Column: Real mixture showing positive deviation from ideal behavior (for example; ethanoltoluene) 3rd Column: Real mixture showing negative deviation from ideal behavior (for example; acetonechloroform)

yressur e diagram

I

x1 Boiling diagram I

bl

0

1

1

0

u XlYl

0

1

P=const

Equilibrium diagram

1 0

1 0

Cl

P =const.S=const

Activity coefficient diagram

dl

Relative volatility diagram

enthalpy Free excess of

dSE

I loo~oc~n

0

the mixture

fl

1oi

/1

__c

Mixing enthalpy and excess entropy

41

x1

3xonst,p=const

O M T-A;, 0

T*AZE

dii

41

1.4 Phase Equilibria

according to the requirement that the system is at thermal equilibrium. In Fig. 1-19 the graphical development of the procedure is shown. This is generally valid and includes real mixtures. Distillation processes are influenced by the working pressure, as shown in Fig. 1-19. With a reduction in pressure, at the same mole fraction x1 in the liquid, the vapor fraction y , increases; hence the separation of the mixture is more efficient at lower pressure. To describe phase equilibria for real binary mixtures, the equations listed in Table 1-9 can be used. Pressure, boiling, and equilibrium diagrams, and other plots showing system behavior characteristics for a real mixture, are given and discussed in Fig. 1-20. According to Raoult's law, and therefore applying to ideal behavior, the activity coefficient can have positive (yi> 1) and negative (yi < 1) deviations. Pressure, boiling, and equilibrium diagrams for the steam distillation of common and important binary mixtures of partial miscibility are shown in Fig. 1-21. Azeotropism An azeotropic point is characterized by the following : 0

0

Pi

X2,ac = Y2,ac

(1-101)

The relative volatility of an azeotropic mixture a,,= 1, therefore (1-102) hence

-

P2

XI

Ylxi

cl

IP iPl

-

iP1 XI

-

YI XI

p=const

-

p=const

t __

x1oc XI

-----c

1

0

-x1oc

x1

I

Fig. 1-21. Pressure diagram (a), boiling diagram (b) and equilibrium diagram (c) of a binary mixture with a miscibility gap over the complete concentration range (left column) and a binary system with a limited miscibility gap (right column). x ~ Azeotropic , ~ ~ concentration B Bubble point line D Dew point line

The equilibrium composition is the same for both vapor and liquid phases X1,ac = Y l , a c

I

P

0

(for an explanation of Eqs. (1-102) and (1-103), see Chapter 1.5) For an azeotropic mixture the isothermal partial pressure curve p 1(xl),p2(x2),the isobaric boiling line t?(xl) and dew line d ( y J and the isobaric equilibrium curve y 1(x,), exhibit maxima/minima (maximum boiling or bubble point and therefore minimum vapor pressure; minimum boiling or bubble point and therefore maximum vapor pressure see Fig. 1-20)

It is not possible to separate an azeotropic mixture by simple distillation. To

42

1 Basic Concepts

t

s,

X-

Fig. 1-22. Influence of the operating pressure on the azeotropic point. A, Azeotropic point at an operating pressure of

PI

A, Azeotropic point at an operating pressure Pll
make separation possible, the position of the azeotropic point must be adjusted. The following steps are required to change the location of the azeotropic point, or to remove if from the diagram: 0

0

If the azeotropic point is restricted only to a certain pressure or temperature range, an increase or decrease in the working pressure can make the azeotropic point disappear (Fig. 1-22) A selected third component may be added to the original binary mixture, causing the azeotropic point to disappear (extractive distillation, Fig. 1-23). The third component is chosen such that it does not form an azeotropic point with either component of the original mixture. Its boiling point should be significantly higher than the highest boiling point of either component of the original mixture and it must be miscible at process conditions with both components of the original mixture

Fig. 1-23. Influence of an auxiliary component on the azeotropic point in an extractive distillation process. Example system: Acetone-trichloro methane Auxilliary component: Methyl isobutyl ketone EC 1 Equilibrium curve without auxilliary component present EC 2 Equilibrium curve with auxilliary component present

0

A selected third component is added to the binary mixture such that, it forms a low-boiling azeotropic mixture with either one of the components of the original mixture. Separation by distillation is now possible (azeotropic distillation, Fig. 1.24). The boiling point of the third component should be in the same range as the components of the binary mixture. At process conditions, the third component should be miscible with both of the original components

Coexistence of the Vapor and Liquid Phases within a Limited Concentration Range It is possible, particularly for distillation processes at high working pressure, for the critical pressure of either one or both components to be lower than the working pres-

1.4 Phase Equilibria

43

With the requirement that C y i = 1 for the vapor phase, the equation needed to compute the boiling or bubble point of a ternary mixture consisting of the components 1, 2, 3, is

x, . P0,l ~

P

P

P

=

1

(1-105)

With the requirement that Exi = 1 for the liquid phase, for a ternary vapor mixture, the computation of the dew points is analogous.

x-

Fig. 1-24. Influence of an auxilliary component on the azeotropic point in an azeotropic distillation process. Example system: Cyclohexane-benzol Auxilliary component : Ethylacetate EC 1 Equilibrium curve without auxilliary component present EC 2 Equilibrium curve with auxilliary component present

sure. Liquid and vapor phases coexist only in a limited concentration range (or mole fraction of the components), as shown in Fig. 1-25. Separation by distillation is only possible within the range of the coexisting phases. If, for example, gas mixtures are condensed under high pressure, low temperature distillation may occur.

Representation of Vapor-Liquid Equilibrium for Ternary Mixtures

For each individual component in an ideal mixture, Raoult’s and Dalton’s laws are valid Po, i x . y . = -. P

+ X z * - Po,2 + x 3 * -

(1-104)

Y , * - P +Y2’--P + y 3 * - - - - 1 Po, 1 Po,2 Po.3

(1-106)

Boiling (or bubble) and dew point curves for ternary mixtures, may be represented by the equilateral triangle discussed in Chapter 1.4.2.2. At constant temperature, the boiling and dew lines, according to Eqs. (1-105) and (1-106), are straight lines within the equilateral triangle. The positions of the lines may be found by an easy graphical procedure if the boiling diagrams for the three binary pairs are given. If the boiling line (liquid isotherm) and dew line (vapor isotherm) are known, the composition of the vapor phase corresponding to a given composition of the liquid phase may be calculated from Eq. (1-104). The positions of the equilibrium lines in the triangle are determined stepwise. For real mixtures, the correlations listed in Table 1-9 may similarly be applied. The liquid and vapor isotherms are curved lines in the triangle diagram. Depending on the behavior of the mixture, the equilibrium or equilibrium distillation lines form one, two or three distillation areas (see [2.1]).

44

1 Basic Concepts bl

a)

Fig. 1-25. Boiling diagram (a) and equilibrium diagrams (b) for a system showing a limited range of coexistance between the vapor and liquid phases. Example system : Carbon dioxide-sulfur dioxide [2.1]. B Bubble point line D Dew point line p r = 40 bar p l , = 91 bar prlr = 62 bar X, 33%, x, = 72.5% Carbon dioxide (At =96 bar, the points A, and A, are the same point for this system, therefore, the bubble point line and dew point line do not show a wet vapor region) Critical point of carbon dioxide: 73.8 bar Critical point of sulfur dioxide: 78.6 bar x Molar fraction of carbon dioxide in the liquid phase y Molar fraction of carbon dioxide in the vapor phase r9 Temperature

Vapor-Liquid Phase Equilibrium of Multicomponent Mixtures

For a mixture consisting of n components, the calculation of the bubble or boiling points including the equilibrium constant (distribution coefficient) K: (defined by Eq. (1-46)) is n

cxi.KT=1

The equilibrium constant K: may be determined from the correlations listed in Table 1-9. Generally K: is dependent on the pressure, temperature and composition of the mixture. For an ideal mixture K: = p0,Jp,hence Eqs. (1-107) and (1-108) are the general form of Eqs. (1-105) and (1-106). Boiling and dew points can be determined iteratively.

(1-107)

1

1.4.3.3 Henry’s Law, Gas Solubility

Similarly, the dew points give

I-= yi

1

Ki*

1

(1-108)

The partial separation of a gas component i (absorptive i) from a gas mixture is usually done using a liquid solvent (absorbent). The

1.4 Phase Equilibria

operating temperature is often higher than the critical temperature of the gas component. During the absorption process the gas component is dissolved or absorbed as a gas without condensation taking place. Under absorption conditions, its liquid state does not exist and the measurement of a saturated vapor pressure is impossible. Describing, for example, the vapor-liquid equilibrium by the expanded Raoult’s law p i = po,i yi. x i , would lead to incorrect results. To find a concept to describe the technical importance of gas solubility, the mixture is assumed to be infinitely dilute, or “an ideally diluted” solution. With a large excess of solvent, when the gas component i is dissolved, the mole fraction xi in the solution tends to 0 (x,-+O). At vapor-liquid equilibrium, Henry’s law for an ideally diluted solution may be applied to the gas solubility of component i in the solvent pi= Hi xi

Lp = const

(1-109)

In an ideally diluted solution, the partial pressure p i of the gas component i is proportional to the mole fraction xi in the solution. The constant of proportionality is Henry’s constant Hi, which is dependent on temperature, pressure, and the respective excess component. Van-der-Waals forces in the solvent bond the gas component i in the solution; the gas component i is physically dissolved. Henry’s and Raoult’s laws are only valid under limited conditions (see Fig. 1-26). Raoult’s law gives an asymptote to the real partial pressure curve p i (xi) where xi 1. Henry’s law gives the tangent of the partial pressure curve at the limit xi -+ 0. Henry’s law is only really valid for gases which are sparingly soluble. A small mole fraction of dissolved gas i leads to a considerable deviation from the linear relationship between the partial pressure p , and the mole fraction xi. For this reason xi is cor+

3 = const

t

/

/

/

/

I’, /

45

KH,I

-HL

Q‘ /

/ Poi

qFig. 1-26. Validity range of Raoult’s law and Henry’s law. pi(xi)Partial pressure curve of i in a real mixture RL Asymptote to the partial pressure curve as xi-+ 1, according to Raoult’s law HL Tangent to the partial pressure curve as xi+ 0 according to Henry’s law K H , i Henry constant Hi pi Partial pressure of component i in the gas phase xi Molar fraction of component iin the liquid phase

rected by the marginal activity coefficient yi,, to give

A comparison of this equation with the corrected Raoult’s law equation gives (1-111)

From this, with the given marginal activity coefficient yi,, and an extrapolated saturated vapor pressure po,;, the Henry’s constant Hi may be estimated. (The dependence of Henry’s constant on pressure is given by

46

1 Basic Concepts

(1-112) where q,Pis the partial molar volume of the dissolved gas i in the liquid phase at infinite dilution.) Normally, the solubility of the gas increases with decreasing operating temperature and increasing pressure (Fig. 1-27). For several gases dissolved in one solute, the solubility increases with increasing boiling temperature. As well as Henry's constant Hi, a measure of the capacity of a liquid to dissolve a gas, are the Bunsen absorption coefficient aBu,iand the Oswald absorption coefficient aos,i,which can be found in tables. The Bunsen absorption coefficient is vNi= a B u , i ' P i

(1-113)

where VN,iis the volume of a gas calculated at standard state, O T , 1.013 bar and dissolved in lm3 of solute. At a partial pressure p i of 1.013 bar of gas i, the Bunsen absorption coefficient aBu,iis the volume at standard state that is dissolved in one unit of solute at a given temperature. The Ostwald absorption coefficient ctOs,; is given by

where cL,;and cG,iare molar concentration of the gas i in the solute (liquid) and gas phase. Therefore the Ostwald absorption coefficient aos,iis the ratio of the concentrations of gas i in the liquid and gas phases. The coefficient is independent of pressure but dependent on temperature. Additional absorption coefficients, such as Kuenen or Raoult absorption coefficients, are listed in [3.1]. These coefficients are seldom used in practice.

tl

6'1

/

xi-

xi-

Xi

+

Fig. 1-27. Sorption isotherms for a gas component i dissolved in a solution in the range of validity of Henry's law. p i Partial pressure of component i in the gas phase x, Molar fraction of component i in the liquid phase yi Molar fraction of component i in the gas phase Molar ratio of component i in the gas phase (loading of carrier gas) XiMolar ratio of component i in the liquid phase (loading of solvent)

1.4 Phase Equilibria

To convert Henry’s constant Hi to the Bunsen absorption coefficient C L ~ , , and ~ the Ostwald absorption coefficient clos,i, the following relationship may be used

47

Eq. (1-116) becomes at absorption equilibrium

-_-.- - Hi 1+y p

Xi l+Xi

L9 = const

(1-117)

where (1-115)

y. ==

moles of component i in the gas phase moles of the remaining inert gas mixture

where Q, is the density and M , the molar mass of the solution. is the loading of i in the gas phase For graphical representation of the gasliquid equilibrium, the “absorption equilibrium”, the partial pressure diagram (pi(xi) moles of component i diagram), and also the equilibrium diagram x.= in the solution yi(xi) or y ( X i ) ,are used. moles of pure solvent If the gas phase is ideal, the solution ideal and diluted, and Henry’s law is valid for the mole fraction of the gas component X i is the loading of i in the liquid phase i, then Figure 1-27 shows the partial pressure lines p i (xi), the equilibrium lines yi (xi) and Hi y.=. x.I L9= const (1-116) y ( X i ) ,the absorption isotherms, at conP stant temperature and physical conditions, (i. e., there are no chemical reactions, and With a solute which is selective to only one the solvent may be described by Henry’s component i of the gas mixture, the solubil- law). Data for practical use giving the solubilities of all the other components of the gas ity of gases in physically and chemically mixture are negligible. During absorption acting solvents, respectively, are found in the remaining gas mixture is inert. Absorp[0.17, 0.21, 1.46-1.561. A method to calcution processes are carried out at low temlate Henry’s constant is given in [1.57]. perature, so vaporization of the solute is Figure 1-28 gives examples of frequently negligible. During the absorption process, the amount of solute and remaining inert used absorption systems. The use of a mixture of pure liquids as a gas are approximately constant. The soluble gas component i in the gas phase may be re- solvent is common in absorption processes. lated to the remaining inert, or insoluble, If, for example, the gas component i is disgas mixture, and in the liquid phase, to the solved in a mixture of two solvents I and 2, solvent. In the gas phase the concentration the Henry’s constant H i , M is, within the of component i is taken as while in the conditions for which it is valid, given by liquid phase it is taken as Xi.Introducing the conversion relationship between the mole fraction and the number of moles of solute in the solvent listed in Table 1-4

48

c 1 Basic Concepts

The solubility of i in the pure solvents 1 and 2, expressed by Henry constant H , , and are required to be known.

i

I

h

1

-

1

-

I

I

1

I

4

Henry's constant for system gas component i-solvent mixture 1 and 2 Henry's constant for system gas component i-solvents 1 and 2, respectively experimentally determined constant (a1,2= 0 for ideal solvent mixture 1 and 2) solvent mixture of 1 and 2 shows a positive deviation from Raoult's law solvent mixture of 1 and 2 shows a negative deviation from Raoult's law mole fraction of solvent 1 and 2 in solution

Gas Solubility with Chemically Active Solvent, Chemisorption

lo-' C2H4

Kr

F Ar

10-2

t0

t I

1 10

I

.J

I

I

1

20

30

I

1

S["Cl

-

I LO

He

50

Fig. 1-28. Solubility of different gases in water. Reference pressure: 1 bar. Representation according to the Linde Report in Jechnik und Wissenschaft". a Absorption coefficient [given in m$t ( m i substance absorbed per t = 1000 kg water)] 19 Temperature

A chemical reaction between the solvent and the gas component i to be separated is called activated absorption or chemisorption. Chemical attractive forces act between the solvent and component i, resulting in high selectivity and solubility. The phase equilibrium becomes a chemical equilibrium. Applying the law of mass action to the chemical equation for chemisorption, the equilibrium constant K: is found. For example, using sodium carbonate solution (Na,CO,) dissolved in water as a solvent for the chemisorption of carbon dioxide (CO,) from a gas mixture, sodium hydrogen carbonate (NaHCO,) is formed according to the chemical equation Na2C0, + CO, 2 NaHCO,

chemisorption

+ H 2 0 ,desorption

(1-119)

1.4 Phase Equilibria

The equilibrium constant KZiis then

tI

49

9 = const

(1-120)

where C N a H C 0 3 ~CNaf203 and cH2C03 are molar concentrations of NaHCO,, Na,CO,, and H,CO,. The maximum solubility of the solvent is limited by the maximum rate of the reaction at equilibrium. According to the Van’tHoff relationships, the equilibrium constant KZi depends on pressure and temperature. Generally, chemisorption processes are more effective at lower temperature and higher pressure. Compared with physical absorption, chemisorption has the following advantages : 0

0

Higher selectivity of the solvent for the separating gas component Higher absorptivity of the solvent and higher rate of absorption

A disadvantage of chemisorption is the fact that the recovery of the solvent might not be effective or even not possible via a reversible reaction. As shown in Fig. 1-29, the positions of the partial pressure lines (absorption isotherms) pi(xi) are different for chemical and physical absorption due to the marginal laws for the partial pressure at pi(xi-,0). According to HAASE[1.87] at the zero point (pi 0, xi 0) the following condition for the partial pressure line pi(xi)applies -+

XI

Fig. 1-29. Partial pressure curves for physical and chemical absorption up < 1 Chemical absorption with association up = 1 Physical absorption up > 1 Chemical absorption with dissociation Partial pressure of the absorbate in the pi gas phase Molar fraction of the absorbed compox, nent i in the liquid phase

where

up number of molecules on the right-hand

K

side of the reaction equation for chemical absorption (in Eq. (1-119) up = 2), if one molecule of sorbate i is dissolved constant (similar to Henry’s constant Hi)

The following cases must be distinguished (see Fig. 1-29): 0

+

0

0

(1-121)

-

up > 1 Chemical absorption including dissociation, a horizontal tangent may be drawn on the partial pressure line at xi = 0. This is a com-

mon case for chemical absorption processes up = 1 Physical absorption, Henry’s law is valid, neither dissociation nor association up < 1 Chemical absorption with association, a vertical tangent may be drawn on the partial pressure line at xi = 0

50

1 Basic Concepts

Fig. 1-30. Determination of the equilibrium curve for adiabatic absorption. a) Balance schematic AB Absorber b) Course of equilibrium curve EC, AB Adiabatic adsorption equilibrium RS Regenerated solvent curve LS Loaded solvent t9,,&, . .. Absorption isotherms FG Feed gas mixture Purified gas mixture, reduced in PG absorbate

Whilst in physical absorption processes the partial pressure p i of the absorbed component i increases approximately proportional to the mole fraction xi in the solution, during chemical absorption when xi is small p i is initially very low (good initial loading capacity of the solvent is the basis of the case up > 1). With increasing xi,p i shows a sharp increase. Absorption is an exothermic process, i. e., during the absorption process heat is released. If this heat cannot be removed, the temperature of the phase streams increases. Since the residence time in the absorber for the gas phase is short and its heat capacity is small at low absorption pressure, the heat of absorption is mainly removed by the liquid phase. In adiabatic absorption processes a significant increase in temperature

of the liquid phase occurs with the consequence of a decreasing loading capacity for the absorptive (absorbant). The course of the equilibrium curve or the operating line for adiabatic absorption at different isothermal states, may be derived using enthalpy balances. For 1 kmol of solvent, the enthalpy balance at constant gas phase temperature, using the notation in Fig. 1-30, is

It follows that the temperature 19 for each kmol solvent loaded with X kmol of absorbed component, is

1.4 Phase Equilibria

Fp, and Cp, are the molar heat capacities of the solvent and absorbed component, respectively, AEAb is the absorption enthalpy (see Chapter 3.3) and dg and 19,are the temperatures of the gas and liquid phase, respectively. For given values of Xi, X 2 , . . the corresponding temperatures I9,, I9, . . . may be calculated using Eq. (1-123). The equilibrium curve for an adiabatic process in the absorber is formed by the intersections of the ordinate parallel lines at Xi, X,, and the corresponding absorption isotherms IP,, d2 (see Fig. 1-30).

tI

a1

.

‘Solubility limit I

/

I

1.4.3.4 Boiling Equilibrium of a Solid Solution, Decrease of Vapor Pressure and Increase of Boiling Point A solid solution consists of surplus solvent and physically soluted, molecularly dispersed, distributed solid. The partial pressure of the solid is generally negligible. Only the pure liquid, in this case the solvent, of the solid solution is vaporized. As the amount of solvent decreases, the solid fraction increases. The boiling equilibrium, (the phase equilibrium of boiling solution and vapor), is of practical interest for the concentration of a solution by vaporization of the solvent. The boiling equilibrium of a solution is characterized by the vapor pressure line p , ( T ) (see Fig. 1-31); this is approximately parallel to the vapor pressure curve po,l( T ) of the pure solvent and may be derived from the equilibrium condition of Eq. (1-49)

(1-124) where ai is the activity of the solvent in the solution and and VII are the molar volumes of the solvent in the liquid and vapor phase.

vI

51

I

v d AT L l n c r e a s e

in bubble point

T-

Fig. 1-31. Vapor pressure curves for a solution and solvent. Mole fraction of dissolved substance x, x2, I < x2,2, etc. Vapor pressure of the solution pL l/T Reciprocal temperature Ap Vapor pressure reduction AT Increase in bubble point P ~ , ~ ( Vapor T) pressure curve of the pure solvent pL Vapor pressure curve of the solution pL,po,l Vapor pressure of the solution and the solvent, respectively Tr Triple point

(n

Integration between the limits of the vapor pressure of the solvent po,iand the vapor pressure of the solution p L gives PL = Po.l

*

a1

(1-125)

or for ideally dilute solutions Raoult’s law PL = P0,l ‘ x1

(1-126)

The vapor pressure p L of the solution is always lower than the vapor pressure of the pure component when they are both at the same temperature. From Eq. (1-126) the reduction of the vapor pressure pa, - pL is

52

1 Basic Concepts

where the expression within the brackets is the ebullioscopic constant kE of the solvent. Ebullioscopic constants for frequently used solvents are listed in Table 1-10. The ebullioscopic constant gives the increase in boiling temperature of the solution if 1 mol If the soluted substances are dissociated in of solid is dissolved in 1 kg solvent. This the solution, the number of moles n2 must may be seen from the dimensions of the be replaced by the number of ions n; in constant kE. Eq. (1-127) For practical use the increase of the boiling point may be quickly and easily approx[l-1281 imated using the nomogram given in [0.17]. n; = n 2 . [I + a . (i - l)]

proportional to the mole fraction x2 (x2= 1 - x l ) of the soluted substance

where a is the degree of dissociation, and i is the number of ions in which 1 mol of soluted substance is dissociated. For example potassium chloride dissociates in water KCl -+K+ + C1-, therefore a = 1 and i = 2. A reduction in the vapor pressure pO,,- p L at constant temperature, corresponds to an increase in the boiling point AT at constant pressure, hence the boiling or bubble point of the solution is higher than that of the pure solvent. For a diluted solution the increase in boiling point A T is

Table 1-10. Ebullioscopic constants for different solvents. Solvent

Boiling point at 1.013 bar ( "C)

Ebullioscopic constant kE (K . kg/mol)

Water Ethanol Diethylether Benzol Acetone Chloroform

100 78.4 34.6 80.1 56.2 61.3

0.51 1.22 2.02 2.53 1.71 3.63

1.4.4 Gas-Solid Phase Equilibrium where T, and q,lare the boiling temperatures of the solution and solvent, respectively, is the vaporization enthalpy of the solvent and x2 is the mole fraction of the dissolved substance (Fig. 1-31). The increase in the boiling temperature is directly proportional to the mole fraction of the substance dissolved in the solution. Instead of using the mole fraction x2 in Eq. (1-129) the molality c, = n2/m, for a dilute solution ( n , B n,) is approximately

1.4.4.1 Gas-Solid Phase Equilibrium, Sublimation

Sublimation is the direct phase change of a substance from the solid into the gas state. Desublimation is the reverse process. In thermal separation processes sublimation is of practical interest in freeze or sublimation drying. Heat transfer into the system causes moisture to sublime directly from the frozen solid into the vapor phase (see Chapter 5.10). Desublimation is important for the separation of gas mixtures; if components of the mixture change directly from the gas into the solid state when heat is removed, these components can be sepa-

1.4 Phase Equilibria

rated from the remaining gas mixture (see Chapter 7.4). The gas-solid phase equilibrium at sublimation or desublimation is described by the generalized CLAUSIUS-CLAPEYRON equation, see Eq. (1.71) and Fig. 1-32 (1-131)

sublimes at atmospheric pressure (“simple sublimation”, Table 1-11). For most substances, in the solid state, however, the triple point pressure is considerably lower than 1.013 bar. Sublimation is then only possible in a vacuum or with the aid of a carrier gas. If steam at a constant pressure is cooled below the triple point, a solid is often not crystallized immediately after passing the sublimation pressure line. Crystallization

where dp/dT is the slope of the sublimation pressure curve at temperature T. FI and 6 are the molar volumes of the substance in the gas and solid states, respectively, and Ah,, is the molar sublimation enthalpy. At the triple point this is the sum of the melting enthalpy and vaporization enthalpy Ah,,,. = Ahs,l -k Ahl,g

(1-132)

For the vapor pressure curve the introduced s and = i?. T/p simplifications are also valid for the sublimation pressure curve. Considering Eq. (I-131), an approximation for the sublimation pressure curve is then

vII v,

vII

(1-133) If the vapor pressure for the solid at the triple point is 1.013 bar or more, the solid

T-

Fig. 1-32. p , T-Diagram of a one-component system. Tr Triple point MC Melting curve VC Vapor curve SC Sublimation curve OC Oversaturation curve S Subcooling point [7.1] p Vapor pressure T Temperature

Table 1-11. Sublimation and melting data of substances which easily sublime [7.1]. Substance

Sublimation temperature (“C) at a sublimation pressure of 1.013 bar

Melting temperature (“C)

Melting pressure (bar)

Acetylene Carbon dioxide Uranium hexafluoride Aluminum chloride Ammonium chloride

-83.6 -78.5 56.4 177.8

- 81.8 - 56.6 69.2 190.0 520.0

5.28 2.03 2.54 35.0

337.8

53

1.18

54

1 Basic Concepts

takes place after passing the surfeiting line, a line parallel to the sublimation curve (desublimation delay, Fig. 1-32). This phenomenon is similar to the supersolubility of a solid dissolving in a solvent (see Chapter 1.4.5).

1.4.4.2 Gas-Solid Phase Equilibrium with Adsorption/Desorption and Convective Solid Drying (Adsorption Equilibrium) Solids with an active interface (adsorbent) can hold on to vapor or gas molecules (udsorbed molecules, adsorbate) at their surface (see Chapter 4). In physical adsorption (physiosorption) the binding forces between the adsorbing surface of the solid and the vapor or gas phase molecules are Van-der-

Waals forces - mostly electrostatic or dispersion forces. In chemical adsorption (chemisorption) the adsorbent and the adsorbed substance, (the adsorbed or bonded form of the adsorbate) have a chemical bond due to valency forces between the molecules. If adsorption equilibrium is reached, no further adsorbate is taken up by the adsorbent. Under these operating conditions of pressure and temperature, the adsorbent is saturated with adsorbed substance. The adsorption equilibrium is characterized by the concentration of the adsorbate in the gas phase and the corresponding equilibrium load of adsorbate adsorbed by the solid. Adsorption equilibrium may be represented by three graphical methods; adsorption isotherms, adsorption isobars and adsorption isosteres (Figs. 1-33 and 1-34):

Fig. 1-33. Adsorption isotherms of water on a Baylith T 144 molecular sieve (Na-form, 3 lo-'' m) and silica gel [4.23]. SG Silica gel BA Baylith X i Mole ratio of water in adsorbent p i Partial pressure of steam 9

1.4 Phase Equilibria a)

55

bl

pi= lOmbar

I

pi= 5 mbar

p, = 2 mbar

I

0

50 S[OCI

-

100

0

Fig. 1-34. Adsorption isobars (a) and adsorption isosteres (b) for water on a Baylith T 144 molecular sieve (Na-form, 3 . lo-'* m) [4.23]. X i Mole ratio of water in adsorbent

t9 Temperature

p i Partial pressure of steam

a

0

a

Adsorption Isotherms: The mole ratio of the adsorbed component in the adsorbent X, is plotted against the partial pressure pi of the adsorbate in the gas phase. Another concentration scale may also be used instead of partial pressure. The partial pressure p i may be replaced by the ratio pi/po,,, which is known in drying technology as the relative humidity pi. (p0,,is the saturated vapor pressure of the adsorbate at the reference temperature 8) Adsorption Isobars: The mole ratio of the adsorbate in the adsorbent Xi is plotted against the temperature at a constant partial pressure of the adsorbate pi Adsorption Isosteres: The temperature is plotted against the partial pressure of the adsorbate pi. The parameter is the mole ratio X i of the adsorbed substance in the adsorbent. (A logarithmic plot of the variables may be more useful)

Figure 1-35 shows characteristic plots of some adsorption isotherms, for different pore systems in the adsorbent and their interactions with the adsorbate. The convective drying of damp solids is characterized by adsorption isotherms. The exchanging adsorbent is the solid and the adsorbate (or adsorbed substance) is the moisture, usually water. A solid is termed hygroscopic if the partial pressure pi of the moisture at equilibrium in the gas space above the solids surface is lower than the saturated vapor pressure p0,; of the moisture at the solids surface temperature. With hygroscopic solids p i is dependent on the temperature and moisture content of the solid. Moisture is bonded by adsorption to the solid and is therefore at a lower energy level than the liquid. To dry a hygroscopic product the specific vaporization enthalpy of the moisture and the bonding energy must be supplied. A solid is not hygro-

56

1 Basic Concepts

Fig. 1-35.Characteristic sorption isotherms. Representation according to KNEULE[5.2]. Form 1 Pure adsorption with mono- or multimolecular layer at the adsorber surface Form 2 Adsorption and capillary condensation Form 3 Adsorption and capillary condensation, nonhygroscopic product behavior at higher product moisture content Form 4 Pure adsorption with low product moisture content, followed by nonhygroscopic product behavior Form 5 Unfavorable course of the adsorption isotherm Form 6 Adsorption to p,, at p, hydrate formation (with water), to pz adsorption and then with cp > cpz capillary condensation

adsorbate in the gas mixture is higher than the partial pressure, corresponding to the adsorbent load X i at equilibrium. Desorption occurs if the concentration of the adsorbate in the light phase is lower than the corresponding equilibrium concentration. If rpi is almost equal 1 condensation of adsorbate takes place. If the structure of the solids consists of a capillary system with small pore sizes, even with pi < 1, capillary condensation of adsorbate with wetting behavior occurs. According to the Gibbs-Kelvin equation (1-134)

where 0

surface tension (N/m)

Mi molar mass of the adsorbate (kg/kmol)

el,; density of the adsorbate in the liquid r

state (kg/m3) radius of curvature (m), negative for wetting (concave surface) and positive for nonwetting- (convex surface) in Eq. (1-134) bl

al

scopic, if the partial pressure p , is equal to the saturated vapor pressure of the adsorbate the partial pressure p iis then only a function of temperature, and only the heat of vaporization is necessary for drying. In practice often low concentrations of the gas phase adsorbate must be handled. The adsorption equilibrium curve is therefore characterized by high loading rates of the adsorbate into the adsorbent at low adsorbate partial pressures in the equilibrium gas phase (approximate vertical tangent of the adsorption isotherm at pi 0, Fig. 1-34). Adsorption takes place when the partial pressure p i or the relative saturation rp, of +

PL

-

t

k

:"i

9-

- - - Capillary condensation

Fig. 1-36. Adsorption isotherm (a), sorption hysteresis (b). Ad Adsorption isotherm De Desorption isotherm X Loading of adsorbent with adsorbate p , Equilibrium partial pressure p Relative humidity

1

35

1

30

25

1/x

20 15 10

5

0 I1

0

0.21 I

01

I I

1.o

106

;

x,y-

08

I

I

I

I

1

I

X

0.6

0.4 0.2

02

I

02

I

01

1

06 x.y

08

1

10

I

I

f 0.8

0

0

04

06

Y-

08

I Adsorption equilibrium data for the system oxygen/nitrogen/active carbon. Presented in a triangular diagram using mass fractions as the concentration scales (8 = - 150"C, p = 1 bar). A,B State points of active carbon loaded with pure oxygen and pure nitrogen, respectively. Curve A.. E.. B Locus of the state points of active carbon loaded with gas mixture. RE Conode TI Adsorption system : Acetylene/ethylene/silica gel ( 8 = 25 "C, p = 1 bar). I11 Adsorption system: Acetylene/ethylene/active carbon (8 = 25"C, p = 1 bar). X Loading of the adsorbent by the key component (adsorbate), in kg key component/kg adsorbent x Molar fraction of acetylene in adsorbate y Molar fraction of acetylene in the gas mixture

58

1 Basic Concepts

1 1

X

0

0

1.0

1.0-

t

1 Y-

Fig. 1-38. Influence of pressure on the course of the adsorption isotherm for the coadsorption of two gases [0.8]. X Loading of the adsorbent by the key component x Molar fraction of the component which is more easily adsorbed in the (key component free) adsorbent y Molar fraction of the more easily adsorbed component in the gas mixture

At the concave surface of a wetting adsorbate in the liquid state, for a capillary of radius r, the adsorbate vapor pressure (po,i ) r is lower than the vapor pressure po,i for a flat surface. For a nonwetting adsorbate, b0Jr > capillary condensation does not occur. Due to capillary condensation, with a monomolecular layer at the adsorbed surface the adsorbate concentration is considerably higher than the saturation concentration X,,, (Fig. 1-36). The adsorption isotherm can be found experimentally using volumetric or gravimetric methods. Special methods to measure proportional properties to detect the change in the occupation of the adsorbate on the surface of the adsorbent are also available (see [1.32, 4.1, 4.21). For the correlation and extrapolation of equilibrium data for adsorption processes with one adsorbate component useful approximations are listed in Table 1-12. Experimentation has shown that different adsorption isotherms are occasionally obtained, for data taken during adsorption or desorption under the same conditions. This hysteresis in the sorption isotherm (Fig. 1-36) must be considered in the design of adsorbers and thermal dryers. The hysteresis phenomena can be explained by capillary condensation. For the adsorption in an adsorbent composed of two components from a gas or vapor mixture, the adsorption equilibrium may be conveniently represented using a triangle. The triangular axes are then l/X, y, and y,x mole fractions or mass fractions. Figure 1-37 shows some practical examples from [0.8]. The effect of the operating pressure of the adsorber can be seen in Fig. 1-38. Published adsorption equilibrium data for binary and ternary mixtures can be found in [1.40]. Adsorption equilibria for two component and multicomponent systems may be calculated using the correlations listed in Table 1-13.

1.4 Phase Equilibria

59

Table 1-12. Correlations and methods for the extrapolation of equilibrium data for one component adsorption. 0

Freundlich equation [1.33] X = k, * pk2

k,, k2 Temperature dependent constant determined experimentally. Freundlich's method is valid for the description of adsorption isotherms of form 5 given in Fig. 1-35. 0

Langmuir equation [1.34] X=

k~ . Xmax . p 1+ k A ' p

kA Adsorption coefficient, Xma,maximum monomolecular layer loading of the adsorbent by the key component (adsorbate).

k System specific adsorption constant (usually k = 1); heat of vaporization of absorbate. AfiAd Adsorption enthalpy, Langmuir equation is valid to the point that the monomolecular layer is saturated. kA and Xma,are experimentally determined. 0

Brunauer, Emmet, Teller equation (,,BET" equation) [1.36] kA*p

x = xma, ' -. 1-p

1 -(n+l).p"+n.p"+'

1 + (kA - 1) * p - kA * p n t '

(In the range 0 < p < 0.35 simplification according to X=

k~ . P (1 - a) . (1 - + kA * Xmax.

and rearranging, gives

-.-

-

1

kA

-1

+ .P Xmax*kA X m a x . k ~

X 1-P the linear relationship in the 1/X. p/(p - l), p diagram. Basis for the determination of the BET surface [4.26]).

n Number of molecular layers of the adsorbed component on the adsorbent (for n = 1 Langmuir equation).

BET equation is valid for multilayer adsorption. Modified BET equation also applicable for systems with capillary condensation [1.36].

0

Dubinin equation i1.371

V Adsorbed volume at the adsorbent

V, Adsorbent saturation volume at p = 1

C Structural factor (depends on the pore structure of the adsorbent)

p Affinity constant, m exponent, characteristic for adsorption system

1 Basic Concepts

60

Table 1-13. Calculation of the adsorption equilibrium of binary and multicomponent systems. 0

Modified Langmuir equation [1.40]:

1

e Experimentally determined adaption parameter X,,,,, Monomolecular layer of component i or of all m components (no uniform fixed value)

k, Adsorption coefficient of the ith component (determined by the adsorption isotherm of i ) 0 0

Modified Dubinin equation [1.41] Ideal Adsorbed Solution Theory (IAST) of MYERSand PRAUSNITZ [1.42], Real Adsorbed Solution Theory o f COSTA et al. [1.43].

1.4.5 Liquid-Solid Phase Equilibrium 1.4.5.1 Solubility of Solids in Liquid Solvents If a solid with an amorphous or crystalline structure is dissolved in a large excess of solvent, the dilute solution gives one homogeneous phase. If more solid is added to the diluted solution, up to the point where the solid can no longer be dissolved, the solution is saturated or at solubility equilibrium. The maximum solubility or the maximum capacity of the solvent to dissolve the solid has then been reached. The solubility of the substance to be dissolved in the solvent depends on the temperature, molar amount of solvent in the crystals, and hence the form of the crystals; the dependence on pressure is only minor. For the same crystal form, solubility usually increases with a higher temperature. If the molar volume of the solid decreases during

dissolution, increasing the pressure also increases the solubility. Figure 1-39 shows a schematic section of the phase diagram for a binary system (solvent - dissolved substance), including the eutecticum. The eutectic of solutions diluted with water is called the cryohydrate. It appears in all solvent - dissolved substance systems. In Fig. 1-39, cooling a solution from point A down to the solubility curve SC, causes crystallization to take place (the case of over solubility will be discussed later). The mass fraction of solvent in the solution thus increases because of partial crystallization of the dissolved solid. During the cooling process, the change of state of the solution moves along the solubility curve SC until the eutectic point E is reached, the solution then freezes completely. If a solution is cooled from B, then the solvent freezes when the ice curve IC is reached. The dissolved substance becomes enriched (the concentration of the solution is increased because the fraction of solvent

1.4 Phase Equilibria

I"C1

W-

Fig. 1-39. Phase diagram for a binary solvent/ dissolved substance system with eutectic point without solvate formation. I Homogeneous solution region I1 Two phase region, solution, and solvent as the bottom layer 111 Two phase region, solution, and solid as the bottom layer I V Region of solid phase, solvent, and eutectic point V Region of solid phase, solute, and solvent IC Ice curve SC Solubility curve E Eutectic point (cryohydrate) ~5 Temperature w Mass fraction of the dissolved substance in the solution Phase. -+ Change of the solution state by cooling from A and B

is decreased by freezing the solvent). During the cooling process, the change of state of the solution moves along the ice curve IC until the eutectic point E is reached; the solution then solidifies. No further separation of solvent or dissolved solid by cooling is possible while passing through the eutectic point E. Therefore, the eutectic temperature is the limit for crystallization by cooling. The closer the

61

system is to the eutectic point the more difficult it becomes to separate the dissolved substance from the solvent. The effect of temperature on the solubility is shown by the solubility curve SC. In crystallization processes, the solubility is expressed as a function of the loading X (kg/kg) of the substance dissolved in the solvent. In the solubility-temperature diagram, the saturation concentration or solubility X i s plotted against the temperature r9, with the result being the theoretical solubility curve or saturation line X ( v )(Fig. 1-40). In Fig. 1-41, if a solution is cooled carefully without a crystallization seed, the following is observed: despite the fact that the saturation line has been crossed no dissolved solid crystallizes. After passing the first oversolubility curve, single crystal seeds may be observed, whilst on passing the second oversolubility curve, small crystals develop spontaneously. The area of the stable, unsaturated solution is followed by the area of the metastable, over-saturated solution between the saturation line and the second over-saturation line. It follows that the area where seeds are spontaneously formed is the area of unstable two- or multiphase systems. The range of practical crystallization processes carried out inside the important metastable area depends upon many process parameters, including solution concentration, intensity of mixing in the crystallizer, cooling rate, initial temperature, number of crystal seeds and the crystallizer shape (surface condition). In order to crystallize an over-saturated solution within the metastable area, inoculate crystals must be present when the solubility curve is crossed. If the crystal seeds are present when the metastable area is entered, the over-saturation of the solution will be reduced; this is because the over-saturation is mainly used by the crystal seeds for growth although a few new crystal seeds are formed. Both crystal growth and seed formation occur and

62

1 Basic Concepts

f

2.

t

%

Ig/lOOgl

Region of unstable solution // 1.

oc oc

k Miers-region) /'

Region of stoble solution (subsatured solution)

Fig. 1-41. Temperature-solubility diagram. SC Theoretical solubility curve (saturation line) 1. O C First super solubility curve 2. OC Second super solubility curve X Solubility r9 Temperature

are dependent upon the over-saturation of the solution (see Chapter 7.2.3). For practical use of crystallization processes, data and references regarding the solubility of solids in solvents can be found in [0.17, 1.46-1.50, 7.21. Calculations of the effect of temperature on solubility can be found in [1.111].

tr-l-l-.-pI \ 2

'0

20

LO

60

80 [ O C ] 100

9-

Fig. 1-40. Temperature-solubility diagram, solubility curves for different salts in water. Representation according to PERRY[0.17]. X Loading of solvent with dissolved substance, g of water-free substance in 100 g water (gA00 g H,O), solubility 6 Temperature

1.4.5.2 Melting Pressure Curve The solid-liquid phase equilibrium of a pure substance is described in the state diagram (see Fig. 1-15) by the melting pressure curve p ( T ) . This curve is formed by the connection of state points, where the liquid and solid phase coexist at phase equilibrium. For crystallization processes which start from the melting point, a knowledge of this curve is important. The melting pressure curve may be given based on the general form of the CLAUSIUSCLAPEYRON equation in a similar manner to the vapor pressure and sublimation curves (Eq. (1-71))

1.4 Phase Equilibria

A&, 1 dP dT T . -

(1-135)

(6 V,)

--

where dp/dT is the slope of the melting pressure curve, Ahs, the melting enthalpy, dependent upon temperature and pressure, and the molar volumes in the liquid and solid state, respectively and T the equilibrium temperature. The molar volumes and are almost the same. It is not allowed to simplify Eq. (1-135) by neglecting <. The difference of the molar volumes A V is given by

6

<

q

Depending on the sign, two cases for the difference of A V have to be distinguished for the course of the melting pressure curve:

63

1.4.5.3 Decrease in the Freezing Point An increase in the boiling point or a decrease in the freezing point of a solution containing a nonvolatile component is compared to pure solvent caused by a reduction in the vapor pressure. The reduction of the freezing point A T of a solution is T, - I; as shown in Fig. 1-42 where To is the freezing point (or melting point) of the pure solvent. At To the vapor pressure is the same for the liquid and solid phases of the solvent. T, is defined by the intersection A of the vapor pressure curve VC and the sublimation pressure curve SC of the solvent. If only pure solvent freezes, the freezing point T of the solution occurs at the intersection B of the vapor pressure curve of the solution VCS and the sublimation pressure curve of the solvent SC. For a dilute solution containing nondissociated, nonassociated and nonvolatile substances the decrease of freezing point is

Case 1: AV>O, 6 > 5, el <@,. From Eq. (1-135) it follows that dT/dp > 0 and the melting temperature increases with increasing pressure. This is common for many substances, especially metals.

c,

Case2: A l / < O , 5 < el >@,. From Eq. (1-135) it follows that dT/dp < 0 and the melting temperature decreases with increasing pressure. For example, this applies to water, gallium, and bismuth. For some substances the sign of A V changes with increasing pressure, (for example with rubidium, caesium and graphite) the sign changes from plus to minus. A maximum temperature for the melting pressure curve then is when A V = 0 or el = e,. Different approximations for the melting pressure curve are discussed in [1.88].

YGas

Solid phase

St-

phase

T-

Fig. 1-42. Lowering of the solution freezing

point.

- T Lowering of freezing point

VC

VCS

SC p

T

Solvent vapor pressure curve Solution vapor pressure curve Solvent sublimation pressure curve Pressure Temperature

64

1 Basic Concepts (1-1 3 7)

where A & , is the melting enthalpy of the pure solvent and x, is the mole fraction of the soluted substance in the solution. Therefore, for an ideal solution, the decrease of the freezing point is proportional to the mole fraction of the solute in the solution. As shown by Eq. (1-128), for solutions, electrolytes, in which dissociation of the solute occurs, the number of molecules n2

t

h

Crystalline phase

t.-i

A

h

x-

B

x-

A

nlB

B+M

AIB-M A/B-M +BlA-M

AiB-M B/A.M A/B M - W A - M

~~

A

x-

0

A

x----t

of solute has to be replaced by the number of ions n;. Replacing the mole fraction x2 of the solute by its molality c, = n2/m,, in Eq. (1-137) gives

=

c,

a

c,

where C, is the cryoscopic constant of the solvent, and is defined within the brackets. C, is the decrease of the freezing point of a

Fig. 1-43. Melting point diagram of some binary systems. Type I Miscibility in the liquid and solid state of components A and B, mixed crystal formation over the entire concentration range. Type I1 Miscibility in the liquid state, immiscible in the solid state, melting point B diagram exhibits eutectic point. Type I11 Miscibility in the liquid state, partial miscibility in the solid state, melting point diagram shows eutectic point and mixed crystal region. Type IV Miscibility in the liquid state, partial miscibility in the solid state, melting L+B/A-M point diagram with peritectic point and mixed crystal region. Miscibility B/A gap reaches temperature region in which mixed crystals are formed. B Type V Compound formation by the components, immiscibility of the solid phases, development of two eutectic points, E, and E,. L Melt Mixed crystals M A/B-M A/B-Mixed crystal (excess of A) B/A-M B/A-Mixed crystal (excess of B) Eutectic point E Peritectic point P Connection A/B VAB Liquidus line LL Solidus line s

65

1.4 Phase Equilibria

solution when 1 kmol of the solute is dissolved in 1 kg of solvent. Table 1-14 shows the cryoscopic constants C, for some solvents. Table 1-14. Cryoscopic constants of several solvents. Solvent

Water Benzol Bromoform Cyclohexane Nitrobenzene Naphthalene Acetic acid

Melting point at 1.013 bar (“C)

Cryoscopic constant C, (K . kg/mol)

0 5.5 7.8 6.5 5.7 80 17

1.86 5.12 14.4 20 6.9 6.8 3.9

In an ideal solution, the increase in the boiling point and the decrease in the freezing point for a chosen solvent is only dependent on the concentration of the dissolved substance, and not the type (colligative properties). The increase in the boiling point and decrease in the freezing point may be used to calculate the molar mass of a soluted substance. 1.4.5.4 State Diagrams of Binary Systems

for Solid and Liquid Phase Equilibrium

For the basic calculation of crystallization starting from a melt, the state diagrams (melting diagrams, 7;x-diagrams) for the partial separation of binary mixtures must be known. A thermal analysis is carried out by the experimental detection of cooling curves. The temperature during cooling is measured as a function of time in solutions of different composition (see, for example, [1.89]). Methods for calculating crystallization equilibrium are presented in [1.112].

Figure 1-43 shows characteristic diagrams for binary systems with explanations given in the caption.

1.4*6 Enthalpy

Of

Phase Changes

For an isobaric phase transition of a substance, the enthalpy change is usually considerable. If a substance is exchanged between two or more phases in a heterogeneous system, a considerable positive or negative heat of evolution is expected. This is because an exchange of substances between phases always has a n associated exchange of heat. The enthalpy change of a substance associated with the isobaric transition from one phase to another at equilibrium may be described analogously to the CLAUSIUSCLAPEYRON equation as dlnZi - Ahi

~

dT

-

_

l?- T 2

_

(1-139)

where Zi is a characteristic quantity for the phase change and is dependent on temperature, the type of substance, and often pressure and concentration. This is explained for important thermal separation processes in Table 1-15. Usually Zi(7‘) is the phase equilibrium curve, or the line that divides the area at the different states of the phases. If, for example, in a one component system, the vapor pressure Zi is equal to the saturation vapor pressure po,i of substance i in the liquid phase, p J T ) is the vapor pressure curve of substance i. Areas of liquid and vapor (gas) phases are separated by the vapor pressure curve; single points on it represent phase equilibria (see Fig. 1-15). If a chemical reaction takes place, in separation apparatus at constant pressure, Z ( 7‘) is the chemical, temperature dependent, equilibrium constant. Eq. (1-139) is

5. Variable Zicharacterizing the phase change and phase change enthalpy Ahi.

nge of substance i between phase 1 and 2 Phase 2 Number of system components 1 S

3, n

1

2, n

g

2, n

g

1

g

2, n

g

1

g

2

1

1 2, n

1 1

Characteristic phase change variable Z j

Phase separation curve (phase equilibrium curve)

Phase c enthalpy

Thermal process

Recrystallization Melting Solidification (crystallization from a melt) Dissolving Liquid-liquid extraction Crystallization from a solution Sublimation Desublimation Adsorption Desorption Drying Evaporation Condensation Distillation Partial condensation Absorption Desorption Liquid-liquid extraction

Recrystallization pressure Melting pressure

Recrystallization pressure curve Melting pressure curve

Recrysta enthalpy Melting

Solidus line Liquids line

Vapor pressure curve

Vapor pressure

Sublima

Sublimation pressure curve Adsorption isotherm Desorption isotherm

Gas load of the solid Moisture load of the solid

Differen solubilit

Saturation curve (solubility curve)

Solid load of the solvent

Sublimation pressure

Partial vapor pressure Henry constant absorption coefficient Gas load of the solvent Distribution equilibrium constant

Bubble point line Dew point line Absorption isotherm

Sorption isotherm

Adsorpti

Evapora Evapora Absorpti

Mixing differenc

tions: s solid phase, 1 liquid phase, g gas phase.

67

1.6 Minimum Separation Work

then the Van-Hoff’s isobar, with reaction enthalpy Ahi. An average value of the differential enthalpy Ahi of the phase change may be assumed, by integrating Eq. (1-139) in the temperature interval T, q.This may be done, if no measured data is available, and the shape of the equilibrium Zi(7‘) is known, using +

Ah, =

R e T, -T2.(lnZi,r,-1nZi,r,)

G-T,

(1-140)

1.5 Separation Factor and Relative Volatility The separation factor a is generally defined as

The separation factor is therefore a direct measure of the separation efficiency of a separation unit or the whole process, and is thus of practical importance. It is usually dependent on pressure, temperature, and phase composition. Large values for the separation factor characterize a process whith a low separation effort. The more a approaches unity, the more difficult the separation. When a = 1 separation is impossible. The separation factor a is the relative volatility in distillation processes. The light phase is the vapor phase and the heavy phase is the liquid phase; the component with the lower boiling point is the “light” component. Taking Dalton’s and Raoult’s laws into account, Eq. (1-142), the relative volatility of the lighter component 1 referred to the heavier component 2, becomes

a1,2 =

Y1 ‘P0,l ~

.

Y2 P0,2

(1-141)

= f (p,

XI,. ..)

(1-144)

For ideal mixtures, Eq. (1-144) becomes where X, is the mole ratio of the key component leaving the separation unit or separation apparatus in the heavy phase, and Y, is the mole ratio of key component leaving the separation unit or separation apparatus in the light phase. Substituting the mole concentration Xl and Y, with the mole fraction x1 and y1 (see Table 1-4) the separation factor a,,2is a1.2 =

Y1 ‘ X 2 ~

Y2 +

x1

(1-142)

where 1 represents the key component and 2 represents the reference component. For a binary mixture Eq. 1-142 becomes (1-143)

a1,2 =

P0,l ~

Po,2

=f(p, T )

(1-145)

In this case, a1,2is dependent only on pressure and temperature and not on the composition of the mixture. A reduction of the working pressure p in distillation processes causes an increase in the relative volatility a and the separation becomes more efficient.

1.6 Minimum Separation Work The mixing of pure substances increases the entropy. The entropy of the mixture is the mixing entropy ASM plus the sum of the

68

1 Basic Concepts

entropies of the pure substances before mixing. Since the entropy increases, the free energy (free internal energy AF, in isochoric processes, free enthalpy AG, in isobaric processes) decreases. Thermal separation operations are isobaric apart from pressure drops which occur in the separation apparatus. During mixing the change of the free enthalpy ACM of the system, the free mixing enthalpy is equal and opposite to the minimum separation work Wmin required to separate the mixture into its pure components. Wmi,

=

(1-146)

-AGM

When k components are mixed to obtain 1 mol of a real mixture, as discussed in Chapter 1.4.1, the free mixing enthalpy AgM, is k

agM= R . T C xi.h a i < 0

(1-147)

i=1

AgM is negative for the entire concentration range, as seen in Eq. (1-147). The minimum separation work required to separate 1 mol of the k component mixture is k

qmin= - E . TCxi.lnai>O

(1-148)

i= 1

Work must be supplied to the system thus values of Wmin are positive. The actual work supplied to separate the mixture is usually higher than the minimum separation work, as given by Eq. (1-148). Additional energy is required to create additional phases if necessary, to divide phases using mechanical means, to mix or to disperse. The energy losses of the separation units, and the pump work necessary to transport the liquid are not considered in Wmin.

For a given mixing enthalpy AH,, the mixing entropy ASM may be derived from the definition of the free mixing enthalpy

(In ideal mixtures the mixing enthalpy AH, = 0; therefore ASM = - AG,/T and therefore ATM = - R . C xi lnxi > 0.) +

1.7 Mass Transfer Fundamentals The basic principles of mass transfer are discussed in detail in [1.95-1.971. Thermal separation processes are actually mass transfer processes; matter is transported between phases and across phase interfaces. Mass transfer is caused by differences in concentration within a phase and by disturbances of the phase equilibrium. The time taken to return to the phase equilibrium depends mainly on mass transfer, but also on heat transfer (heat is transported not only by convection and radiation at higher temperature, but also by mass). For the design of thermal separation processes, along with a knowledge of phase equilibria, it is also important to have a detailed understanding of how equilibrium is reached and the time required, taking into account restrictions in the mass transfer rate. There are two ways in which matter may be transported by the concentration gradient (the driving force):

Molecular Diffusion: Transportation of molecular size matter. Molecular diffusion takes place in solids and phases with no motion or in phase boundaries Convective Diffusion or Convection : Matter is transported in groups of molecules with the concentration gradient as the driving force, along with the free or forced flow. Under the operating condi-

1.7 Mass Transfer Fundamentals

tions chosen in thermal separation processes, convective mass transfer always occurs in liquid and gas phases Heat and mass transfer are analogous processes. Molecular diffusion in homogeneous materials or phases is similar to heat transfer. Convective diffusion or convection in homogeneous materials or phases corresponds to heat transfer by convection. Mass transfer at the phase boundary corresponds to heat conduction. Mass transfer between phases occurs like heat transfer in several chronological steps. The slowest step controls the rate of the entire process. Thus the mathematical descriptions of heat and mass transfer operations are analogous. Calculation methods and approaches to calculate the heat transfer coefficients may similarly be used to calculate mass transfer coefficients. (See Table 1-18 in Chapter 1.7.2 for the analogy of heat and mass transfer.)

1.7.1 Mass Transfer by Molecular Diffusion 1.7.1.1 Steady-State Diffusion Analogous to Newton’s law of momentum transport and Fourier’s law of heat transfer by conduction, Fick’s first law for mass transfer by steady-state equimolar diffusion, is

x

Dj

69

x coordinate of the diffusion space diffusion coefficient of component i in the diffusion space (m2/h)

Introducing the partial pressure of gases as measure of concentration, Fick’s first law becomes

oi.A n 4,x = - _ . R.T or m I., x =

-

apj

oi-M ~A. R.T

(kmol/h)

ax

api *

~

ax

(kg/h)

(1-150)

(1-151)

Fick’s first law describes equimolar diffusion, in which all components of the system may diffuse independent from each other. During thermal separation processes, matter is transported through phase boundaries. If a phase boundary is selectively permeable to one component, only one-directional diffusion is possible (an especially important case for absorption, adsorption, and drying). For one-directional diffusion, STEFAN’S law gives (1-152)

where c is the sum of the molar concentration of all components in the observed phase, and i is the diffusing component. For gases, Eq. (1-152) becomes

where rate of mass flux of substance i in direction x perpendicular to the area A (kmollh) diffusion area (m2) A aci/ax concentration gradient in the direction of the diffusion flux (kmol/ m3/m) Hi,x

where p is the system pressure. The difference in the flow rate fij,x as calculated by Eqs. (1-149) and (1-152), is the factor c/(c - cj), which is due to the additional superseding one-directional diffusion (“Stefan flux”). The amount of mass flux transferred by diffusion is therefore larger

70

1 Basic Concepts

with one-directional diffusion than with equimolar diffusion.

only for simple geometric cases (for example, plate, cylinder, sphere).

1.7.1.2 Unsteady-State Diffusion

1.7.1.3 Diffusion Coefficient

In unsteady-state diffusion processes, the concentration distribution (or concentration gradient) changes with time and position. Fick's second law for unsteady-state diffusion is analogous to the Fourier equation for unsteady heat transfer

The diffusion coefficient D generally depends upon temperature, pressure, the concentrations of the components and the substance mixture components to be diffused. Diffusion coefficients for several systems are listed (see, for example, [1.47, 1.49, 1.90-1.921) or may be calculated empirically [0.8, 8.1, 8.2, 8.16, 8.171. Diffusion coefficients for some systems are shown in Table 1-16 and simple calculation methods are presented in Table 1-17.

(1-154) Solutions of the partial differential equation for given boundary conditions exist

Table 1-16. Diffusion coefficient for different systems [0.1]. Diffusing component

2"d Mixture component (diffusion medium) (solvent)

Pressure

Temperature

(bar)

("C)

100 300 1095 1249 178 428

Gold

Lead

Silicon

a-Iron

Copper

Silver iodide

Benzol Carbon disulfide Methanol

n-Heptane n-Heptane Water

Benzol

Air

1.013

Benzol

Hydrogen

1.013

Benzol

Carbon dioxide

1.013

0

Steam Steam Steam

Air Hydrogen Carbon dioxide

0.981 0.981 0.981

45 0 0 0

25 25 18 0 45 0 45

Concentration

Diffusion coefficient

(m2/h)

4.5 to 7.1

0.83 . 0.54. 0.54. 1.80. lo-' 0.48. lo-' 1.23. lo-'

50 50 0.25

0.89. lo-' 1.28 * lo-' 0.49. lo-'

0.03 to 0.09

0.0270 0.0364 0.1058 0.1437 0.0189 0.0257 0.083 0.278 0.051

1.7 Mass Transfer Fundamentals

71

Table 1-17. Simple calculation methods for the diffusion coefficient. Diffusion in the gas phase The diffusion coefficient D l , 2of gas 1 into gas 2 under moderate pressure may be approximated using the critical data of the gases according to CHENand OTHMER[8.15]

(1-155) Where M I ,M2 Molar mass, kg/kmol Tk, Critical temperature, K J $ 2 Critical molar volumes of gas components 1 and 2, m3/kmol T Absolute reference temperature, K P Reference pressure, bar Dl,2 Diffusion coefficient, m2/s (To convert to different reference conditions and to therefore compute a rough estimate of the diffusion coefficient at different temperatures and pressures, a diagram is given by SLATTERY and BIRD [S.l5]).

5,

0

Diffusion in the liquid phase In large dilution of a solution without dissociation, the diffusion coefficent D l , 2of component 2 in a solvent 1 may be calculated using an equation given by WILKEand CHANG [8.15]

(1-156) Where D l , 2 Diffusion coefficient, cm2/s T Reference temperature, K q1 Dynamic viscosity of solvent, CP M I Molar mass of solvent, kg/kmol & Molar volume of dissolved substance referred to the boiling point at 1.013 bar, cm3/mol C Association factor (C = 2.6 for water, C = 1.9 for methanol, C = 1.5 for ethanol, C = 1.0 for benzol, ether and heptane as a solvent) The temperature dependency of the diffusion coefficient D l , 2 is approximated by the Stokes-Einstein-Term [0.17] (1-157)

The influence of the concentration on Dl,2is discussed, for example, in [0.17].

72

1 Basic Concepts

1.7.2 Mass Transfer by Convection According to the method used to calculate heat transfer by convection, the convective mass transfer under steady state conditions is

PB

I

t

I

,_

Phase I Ci,G

hi = p; * A * ( q k- Cj,G) = A * Ac;

Y

Phase II

L

(1-158)

X-

Fig. 1-44. Illustration of mass transfer.

See also Fig. 1-44, where: convective flow rate of substance i (kmol/h) area of phase boundary (m2) A cLK,c , concentration ~ of substance i in the bulk of the observed fluid Phase and at the phase boundary, respectively (kmol/m3) Pi mass transfer coefficient (m/h) ti;

PB Phase boundary 6 Boundary layer thickness c, Concentration of key component x Space coordinate

matter is only transported by molecular diffusion. Eqs. (1 - 149) and (1 - 158) are then

For gases, Eq. (1-158), gives

ti = I

P;-A -*

R-T

(pi,k

- P I ,G)

=

P,.A

Re T

. Apl (1-159)

where the partial pressures of the gas i, pi,K and pi,Gare those in the bulk of the gas phase and at the phase boundary, respectively. Eq. (1-158) describes generally the transition of a substance from within the bulk fluid to a phase boundary, or from a phase boundary into the bulk fluid (for the latter case, Eq. (1-158) must be appropriately modified). The concentration of the component in the kernel c ~ is , assumed ~ to be constant throughout the fluid. A concentration gradient Aci occurs only in the vicinity of the phase boundary (see Fig. 1-44). Furthermore, it is assumed that a laminar current exists in the boundary layer and that

and the mass transfer coefficient is

p.=-Di

'

s

(1-161)

where 6 is the thickness of the boundary layer. From this equation, the dependency of the mass transfer coefficient Pi on the diffusion coefficient 0; and the boundary layer thickness 6 of the fluid flow, may be seen. The laminar boundary layer and turbulent bulk cannot be distinguished exactly, due to the continous transition; the boundary layer thickness 6 is, therefore, a formal complementary variable. The mass transfer coefficient P depends upon fluid mechanics (free flow, forced flow), the physical characteristics of the exchanging matter system and the properties of the substance in the fluid phase, and may

1.7 Mass Transfer Fundamentals

73

Table 1-18. Analogy between mass and heat transfer. Analogous variable

Heat transfer

Mass transfer

Transfer variable

Heat flux Q (kJ/h)

Mole flux ri (kmol/h) Mass flux m (kg/h)

Driving force

Temperature gradient A 8 ("C)

Concentration gradient Aci (kmol/m3) Partial densitiy gradient Aei (kg/m3) Partial pressure gradient Api (bar)

Thermal conductivity

Diffusion coefficient D (m2/h)

Heat transfer coefficient a (W/(m2. K))

/3 (m/h or kg/(h . m2. bar) or

Transport coefficient Transfer coefficient

1 (W/(m. K))

Mass transfer coefficient

kmol/(h . m2 . bar) ( k is used as overall mass transfer coefficient)

Dimensionsless numbers to consider -

Gr, modified Grashoff number (see [1.31])

free flow Grasshoff number

-

Re

forced flow Reynolds number

=

W.1 ~

V

Reynolds number

- flow

Galilei number

F, Fr=-=-

- flow

w2

Fg

(.g Froud number

- two phase flow

Weber number - two phase flow

Eotvo's number -

physical characteristics

v

pr=-=a

V.Cp'Q

1

Prandtl number [ratio of molecular momentum transfer (friction, or viscosity effect) to molecular heat transfer (heat conduction)]

V

sc = -

D

Schmidt number [ratio of molecular momentum transfer (friction or viscosity effect) to molecular mass transfer (diffusion effect)] (continued next page)

74

1 Basic Concepts

Table 1-18. (continued) Analogous variable Dimensionless numbers used to calculate the transfer coefficients

Heat transfer

Nu=-

Mass transfer

a.1

Sh =

I

p.1 ~

D

Nusset number (ratio of total heat transfer over heat conduction alone)

Sherwood number or second Nusset number (ratio of total mass transfer over molecular mass transfer)

Nu = f ( G r , , P r )

Sh

for free flow

for free flow

Nu = f(Re, Pr,

. ..)

=f

(Grs,Sc)

Sh = f(Re, Sc,

... )

for forced flow

for forced flow

Nomenclature : Fg gravitational force, F, inertial force, Fq viscosity force, F, surface force, I = L , characteristic length, q dynamic viscosity, v kinematic viscosity, I heat conductivity, a temperature conductivity, e fluid density, surface tension of fluid(s), g gravity constant (9.81 m/s2), w flow velocity of fluid, thermal volumetric expansion coefficient of the fluid (for Grashoff number)

known physical characteristics and heat transfer area. In mass transfer, the fluid phases are in contact, and are separated by a phase boundary. The interfacial area depends upon the internals in the apparatus, fluid mechanics, and the properties of the substances making up the phases. Therefore, the dimensions of the surface of the phase boundary are not easily determined. At the phase boundary, equilibrium between the phases is assumed. This equilibrium at the phase boundary must be considered in mass transfer operations.

be calculated in a similar way to the heat transfer coefficient a , by means of dimensionless numbers. The most common dimensionless groups are listed in Table 1-18, including those essential for the calculation of the mass transfer coefficient.

1.7.3 Overall Mass Transfer Overall mass transfer is the transportation of matter through a phase boundary, from one fluid phase into another. Resistance to the mass transfer is analogous to resistance to heat transfer in that it may be considered as being divided into individual resistances; this will be demonstrated later. The main differences between mass transfer and heat transfer are: 0

In heat transfer, the fluid phases are separated by a defined, solid wall with

0

Individual resistances in heat transfer operations may easily be determined by experimentation. In mass transfer operations between two fluid phases, only the total resistance can be found experimentally, and only under complex experimental conditions. During heat transfer the individual resistances are: resistance due to convection in fluid 1, heat conduction

1.7 Mass Transfer Fundamentals

resistance of the wall, and resistance due

to convection in fluid 2. In mass transfer

operations, a mass transfer resistance has to be considered in each phase of the system. Most approaches (two film theory, theory of surface renewal) assume that there is no transfer resistance in the surface of the phase boundary which is not true for every case. For example, surface inhibition occurs at the boundary if active boundary substances are enriched within the boundary. If the substance exists in both phases under different conditions or forms, there may also be surface reactions. The mass transfer may also be considerably influenced by eddies at the boundary (turbulence in the boundary surface layer, Marangoni effect).

75

PB

t u

Fig. 1-45. Illustration of mass transfer, the twofilm theory.

PB Phase boundary (Interfacial area) Concentration of key component x Space coordinate c

where 1.7.3.1 Two Film Theory, Mass Transfer Coefficient and Turbulence Theory

The two film theory [0.4] describes the mass transfer between two adjacent phases. The main resistance occurs in the two boundary layers at either side of the interface. In these laminar boundary layers, matter is only transported by molecular diffusion. With a phase equilibrium at the interface, the interface itself offers no resistance to the mass transfer. Mass transfer is very fast in the bulk of the phase, due to turbulent convection. The concentrations c ~and , cKII ~ are uniform throughout the bulk phase (see Fig. 1-45). As shown in Chapter 1.7.2 and Fig. 1-45, the mass flux from phase I to phase I1 is

ri

mass flux across interface mass transfer coefficient in phase I and phase I1 c ~ , cK,[, ~ , bulk concentration of the substance of interest in the bulk of phase I and phase 11, respectively A interfacial area c ~ cG,II , ~ , concentration of the substance at the interface

P,,PII

The existing phase equilibrium at the interface means that the concentration is (1-164)

CG,I = K * . CGJI

where K* is the equilibrium constant (see Chapter 1.4 and Fig. 1-46). The concentration gradient in phase I is (1-165)

Matter leaving phase I enters phase I1 through the interface and therefore,

and in phase I1 is n

(1-166)

76

1 Basic Concepts

C

Introducing an overall mass transfer coefficient kll, related to the overall mass transfer resistance in phase 11, it follows that 1

1

1

k11

PI1

K * *PI

-=-+-

CII

therefore the molar flux ri is

-

(1-172)

Fig. 1-46. Concentration diagram for mass transfer. BL Balance or operating line

EC Equilibrium curve c1 Concentration of key component in phase I cII Concentration of key component in phase 11

Considering the concentrations cK,I and cK,,, related to the equilibrium concentrations cy and c;"l c K , ~= K " .

~fi

(1-167)

cl*

(1-168)

and

c,,,,

1

=-

K*

(1-171)

Eq. (1-165) may now be written as

where all the variables refer to phase 11. Analogously, referred to phase I 1

1

kl

PI

- _ --

K*

+-

(1-173)

PI1

and

li = k, . A *

(CK,1

- )c:

(1-174)

Eqs. (1-171) and (1-173) Show how the aPpropriate reciprocal value of the mass transfer coefficient is used to calculate the total mass transfer resistance of phases I and 11. When PI1% P, is k, = PI, and the total mass transfer resistance is essentially controlled by the resistance in phase I. When PI % PII is k,, =DI1 and the total mass transfer resistance is essentially controlled by the resistance in phase 11. The ratio (1-175)

combining Eqs. (1-166) and (1-169) gives (1-170) The concentration at the interface cG,1 1 , which is not easily determined experimentally, has been eliminated.

is derived from Eqs. (1-162) and (1-163) and gives the slope of the line connecting PI and Pz in Fig. 1-46. The two film theory is only an approximation of the real mass transfer in thermal separation processes. Nevertheless, it is used for the evaluation of mass transfer measurements. In the thermal separation of mixtures, the interface between phases is usually not

1.8 Steady-State Cocurrent Operation

fixed; it is constantly changed and renewed by the effect of the flow of the phases. More accurate descriptions of mass transfer are shown by HICHBIE[1.93] in his penetration theory and DANCKWERTS [1.94] in his rheory of surface renewal. Both theories take into account the change of the surface of the interface. According to turbulence theory, the mass transfer coefficient /3 is proportional to the square root of the diffusion coefficient D (see Eq. (1-161)).

/3-@

(1-176)

77

(multicomponent) carrier and the component to be exchanged. The total mass balance of the unit (see balance area BAI in Fig. 1-47 and Chapter 1.3) gives the following equation for the exchanged component La * X , - L , . X ,

where . . L a ,L ,

x,, x,

=

Gu . yW - G, . ya (1-177)

entry and exit flux of phase I mole fraction of transferred component at the entry and exit of the unit (phase I) entry and exit flux of phase I1 mole fraction of transferred component at the entry and exit of the unit (phase 11)

1.8 Steady-State Cocurrent Operation

Ga, G, y,, y ,

The cocurrent principle, mentioned in Chapter 1.1, is the basis for the following general discussion of steady-state cocurrent operations. Figure 1-47 shows two immiscible phases Ph I and Ph 11, guided in cocurrent flow through a separation device. Both phases may be mixtures of components, and during the contact of the two phases, one component is transferred from Phase I to Phase 11. Therefore, both phases consist of a

Due to selectivity, only one component is assumed to be transferred between both phases, thus the total flux of each individual phase changes throughout the separation unit. But the flux of the inert components in each phase remains unchanged. Therefore, it is convenient to relate the mole fraction x and y to the inert carrier fluids L, and G, as a mole ratio of component i in

Fig. 1-47. Steady-state cocurrent operation. PhI Phase I PhII Phase I1 BAI Balance area I (Complete separation unit) BAII Balance area I1 (Partial separation apparatus)

78

1 Basic Concepts

the inert (moles i/moles inert). Eq. (1-177) then becomes LT*( X , - Xu) = GT*(Y, - Y,)

tl

(1-178)

In a X X coordinate system, this equation gives a straight line of slope -LT/GT between the points PI (X,, Y,) and P2(X,, Y,) (see Fig. 1-48). Considering only one part of the separation unit (balance area BAII in Fig. 1-47), the mass balance of the exchanged component is L T * ( X , - X ) = G , * ( Y - Y,)

(1-179)

XG

xu x

xu

X----t

Fig. 1-48. Loading diagram for a steady-state cocurrent mass transfer operation from phase I to phase 11. EC Equilibrium curve BL Balance line Y Loading of phase I1 X Loading of phase I

In a K X plot, it also gives a straight line of slope -LT/GT between the points P, (X,, Y,) and P ( X ,Y) (see Fig. 1-48). The line is the balance line, or operating line, of the separation in a steady-state process with cocurrent flow. It is identical to the line given by Eq. (1-178). Points on the balance line represent any chosen cross section of the separation unit, with the corresponding concentration X and Y. P, characterizes the entry cross section into the unit and P2 the exit cross section. (According to Eq. (1-177), if the mole fraction is the concentration scale given by Eq. (1-179), the balance line becomes a curved line. This is also the case if the inert fluxes LT and GT are not constant along x 4 the length of the separation device.) The equilibrium curve Y ( X )of phases I Fig. 1-49. Loading diagram for a steady-state and 11, based on the loading of i, Y and X , cocurrent mass transfer operation from phase I1 is now added to the Y,X diagram in to phase I. Fig. 1-48. The intersection Q(X,, Y,) of EC Equilibrium curve the operating line and the equilibrium curve BL Balance line represents the conditions at the exit cross Y Loading of phase I1 section. This is called a theoretical transfer X Loading of phase I unit, and equilibrium between the phases leaving the transfer unit is reached. If component i is transferred from phase I1 into phase I, the operating line is above ally (for example, cocurrent distillation and the equilibrium line, as shown in Fig. 1-49. cocurrent drying). Countercurrent flow of Cocurrent operation, in thermal separation the phases, which will be discussed in the of homogeneous mixtures is used occasion- following chapter is of more practical use.

tI

1.9 Steady-State Countercurrent Operation

79

1.9 Steady-State Countercurrent Operation a)

Based on the principle of countercurrent flow presented in Chapter 1.1, steady-state countercurrent flow operations are generally discussed in the following section. The separation of a mixture in a single stage does not normally separate the mixture into fractions of the required purity. To increase the separation effect, single stages may be connected to form a cascade (Fig. 1-50). A cascade is a separation device consisting of several similar stages, or several separation units connected in series. If the phases are in countercurrent flow, a serial connection of single stages similar to a countercurrent flow cascade, can be achieved in a countercurrent flow column (Fig. 1-51). This is a practical, simple and economic method of multiplying a single stage separation effect. It is not necessary for countercurrent flow phases to be in stagewise contact (as in tray towers). Constant contact of the phases throughout the length of the column is possible. For the design of countercurrent flow columns, essentially two theories are used,

[ IF M

A

Fig. 1-51. Countercurrent flow column. a) Reflux principle b) Flow through principle F Feed to be separated E Top product A Bottom product I . . .V Separation stages in the countercurrent flow column

the theory of separation stages and the kinetic theory of separation of mixtures in countercurrent flow. The theory applied depends mainly on the type of countercurrent flow process, the internals in the column, the type of mixture to be separated, and its physical and chemical properties.

1.9.1 Theory of Separation Stages

Fig. 1-50. Series connection of individual separation stages (Cascade). I, 11, 111 Equivalent separation stages F. . Feed L , , L,, . . . Individual fraction flows

In separation processes, a common problem is a mixture of a certain composition to be separated into two fractions, where after separation each fraction is of the required composition. If the separation is carried out in a countercurrent flow column, the

80

1

Basic Concepts

necessary height for heat and mass transfer must be determined. The height of the column depends on the number of separation stages connected in series, where each stage represents a single theoretical separation stage (see Chapter 1.1). The actual height of the countercurrent flow column is fixed by the number of theoretical stages and the "stage efficiency factor" (amplification ratio). Determination of the number of required stages is now discussed. Figure 1-52 shows two immiscible phases PhI and PhII flowing countercurrently through a column. Both phases consist of a multicomponent mixture but only one component is transferred from phase 1 to phase I1 during the contact of the phases. Therefore, the phases are composed of a (multicomponent) carrier or solvent and the component to be transferred. Phl PhII

Fig. 1-52. Countercurrent flow column. PhI Phase I PhII Phase I1 BAI Balance area I (total separation column) BAII Balance area I1 (a column section) CC Countercurrent column

Using the notation given for cocurrent flow, a mass balance for the whole separation column (balance area I in Fig. 1-52) for the component to be transferred is

L a .x, + 6,. ya = i, - x, + G, .y ,

(1-180)

where the mole fraction is the concentration scale, or the loading

A mole balance over a section of the column (balance area I1 in Fig. 1-52) with respect to the component to be transferred, gives

and

G,* ( Y - Y,) = i, . (X-X,)

(1-183)

Eqs. (1-181) and (1-183) are plotted on a X X coordinate system and straight lines are produced between the points PI (X,, Y,) and P2(Xa,Y,) and between the points P,(X,, Y,) and P ( X , Y ) .These are the balance lines, or operating lines, with slope LT/GT,the ratio of the flow rates of the carriers. When a component is transferred from phase I into phase 11, the balance line is below the equilibrium line. When a component is transferred from phase I1 into phase I (see Fig. 1-53), the balance line is above the equilibrium line. If the carrier fluxes L, and G , are not constant along the length of the separation device, or other concentration scales are used, the balance line is curved. Points on the balance line link the related mole concentrations X and Y at any cross section of the separation unit. The greater the distance between the balance line and the equilibrium curve, the higher the concentration gradient, i. e. the driving force, for the mass transfer, or the "disturbance from equilibrium". Driving

1.9 Steady-State Countercurrent Operation

Fig. 1-53. Loading, or operating diagram of a countercurrent flow process. BL Balance line EC Equilibrium curve Y Loading of phase I1 X Loading of phase I

forces vary along the column height, as shown in Fig. 1-53. If the balance line touches the equilibrium line, the driving force is zero. The phases in contact at that particular cross-sectional area are in equilibrium, no mass transfer occurs. The separation processes discussed so far consider phase I as a fresh feed into the column. This is important for separation processes such as absorption and extraction. For these processes it is practical to use mole ratio as a concentration scale based on the solvent or inert flux, since the flow remains constant, or will only change negligibly, along the column height. If phase I is generated by phase transition or phase reversal from phase I1 in a heat exchanger on top of a counterflow column, a reflux principle is used instead of the simple flow-through principle (see Fig. 1-51). Phase I is no longer a fresh feed to the column but is the reflux generated by converting phase I1 into phase I; this is the case in countercurrent distillation or rectification. The liquid phase I, the reflux, is generated by partial or total condensation of the

G,y

81

t t

LeX PhII PhI

Fig. 1-54. Reflux principle in a countercurrent flow column.

CC Countercurrent flow column C Condenser BA Balance area PhI Phase I (reflux) PhII Phase I1 (vapor phase)

upflowing vapor phase I1 in a condenser at the top of the column. A mass balance over the upper part of the counterflow column, as shown in Fig. 1-54, at constant vapor flux G and constant reflux with total condensation of phase 11, gives G=L+E

(1-184)

and for the component to be transferred

-

G y

=

i . x + I?.X,

(1-185)

From this, the balance, or operating, line of the rectification column is given by (1-186) where the reflux ratio v is the ratio of the downflowing reflux L = R to the leaving product at the top of the column E

82

1 Basic Concepts

R

(1-187)

V = 7

E

and following from Eq. (1-186)

1.9.2 Method to Determine the Number of Theoretical Separation Stages for a Countercurrent Column

A discussion of a method of determining (1-188) the number of theoretical separation stages v+ 1 required to separate a mixture in a counterflow column follows. Eq. (1-188)is plotted on a y,x diagram (see Fig. 1-55) as a straight line if the slope v/(v + 1) of the balance or operating line is McCabe-Thiele Method constant, or rather if v is constant. A graphical method to determine the theoretical stages of counterflow columns which is easy to use was introduced by MCCABE [1.98]. and BIELE The McCabe-Thiele method is based on the idea of theoretical stages (theoretical separation unit, theoretical tray). From Chapter 1.1, a theoretical stage is that part of a separation apparatus in which heat and/or mass transfer occurs between two phases in contact. Both phases leaving the theoretical separation unit are at phase equilibrium. For a given separation task, the counterFig. 1-55. Operating diagram for a binary system flow column contains the required number EC Equilibrium curve. of theoretical stages in series. To calculate BL Balance or operating line the number of theoretical stages, the x (tan x = v / [ v + I]) Gradient McCabe-Thiele method is applied, for exyo (yo = x E / [ v+ 11) Ordinate intercept ample, to a rectification column for the parx Molar fraction of the low-boiling compotial separation of a binary mixture. In this nent in the liquid phase case, the molar vaporization enthalpies of y Molar fraction of the low-boiling compothe components of the mixture are assumed nent in the vapor phase to be equal and the enthalpy changes of the vapor or the liquid are negligible. Then, the balance line is straight. (A method to calcuWhen, at the phase contact, an equal late a solution to this separation problem amount of matter is converted from phase under different conditions is shown in I (liquid) into phase I1 (vapor) and vice Chapter 2.5.2.4). Fig. 1-56gives a schematic versa the mass fluxes G and L are constant diagram for a rectification column with n along the length of the column. In adiabatic theoretical stages and also the correspondoperation of the column this is only valid if ing y, x diagram according to MCCABEand the components of the mixture have equal BIELE. The vapor flux G leaving the nth stage molar vaporization enthalpies and exhibit with composition yn of the lower-boiling ideal behavior.

y = - . xV+ v+l

XE

1.9 Steady-State Countercurrent Operation

83

Fig. 1-56. Schematic of a rectification

Yn-1

Yn-2

Yn-3

I

I

1 I

&Phase

+Phase

X - - t

I1 (vapor phase) I

(liquid phose. reflux)

component is totally condensed in the condenser. Some of the condensate leaves the top of the column as product E with composition y, = x,. The rest of the condensate, L = G - E, is the reflux at stage n with composition x,. Vapor with composition y , and reflux with composition X , are linked by the balance line given by Eq. (1-188), where in particular y , = x,. The balance line intersects with the y = x line at the point A (xE,yn= xE) in the McCabe-Thiele diagram. For a given reflux ratio v, it follows that yo = x,/(v 1) and therefore point B (O,y,) on the balance line AB is fixed. In general, the balance line combines the mole fraction y,-, and x, of two stages.

+

~

V y,-1 =-. x, + XE v+l v+l ~

(1-189)

Since stage n is a theoretical stage, the vapor leaving the stage yn is in phase equilibrium with the reflux x, leaving stage n. On the McCabeThiele diagram, this gives the point A, with the coordinates x, and y , on

column for the separation of an ideal binary mixture, and the corresponding McCabe-Thiele diagram. C Condenser RC Rectification column n, n - 1,. . . Theoretical separation stage number, indice for the vapor and reflux mole ratio of the low-boiling component EC Equilibrium curve BL Balance or operating line y Molar fraction of the low-boiling component in the vapor x Molar fraction of the low-boiling component in the reflux

the equilibrium curve. The points A,, A,, and A characterize a theoretical stage n. The following stages n - 1, n - 2, etc., can be constructed in the same way, by extending the steps between the balance line and equilibrium curve. For every counterflow separation process, the number of theoretical stages can be found by this simple graphical procedure. Each stage is constructed between the equilibrium curve and balance line, starting with the initial point and finishing at the end point. Obviously in Fig. 1-56 the separation of a binary mixture with the given reflux ratio v gives a top product with composition x, and a bottom product of composition x > xmin. xmi,can only be reached with an infinite number of theoretical stages. Therefore, the reflux ratio v is the minimum reflux ratio for xmi,(see Chapter 2.5.2). If xminshould be reached by the separation, the reflux ratio v has to be raised accordingly (for more details see Chapter 2.5.2). A method similar to the McCabe-Thiele method exists for the graphical determina-

84

1 Basic Concepts

tion of theoretical separation stages for counterflow columns where phase I is not a reflux but a fresh feed phase. The number of theoretical stages is found by the construction of steps between the equilibrium curve and the balance line (see Fig. 1-57). The ratio L T / G T is of particular importance. With practical separation stages installed in counterflow columns, phase equilibrium cannot be reached. Therefore, the efficiency of a practical stage is lower than that of a theoretical stage. For the installation of practical stages in a column, the stage efficiency E (exchange ratio, MURPHREE Efficiency, introduced in Chapter 1.1) is found by a comparison of the number of practical stages and theoretical stages. E may be related to either one of the phases. Relating E to the upflow phase gives

Phl

Egm =

yd yd

- yu - yu,h

(1-190)

or Egm =

yu - yd

'uth - yd

(1-191)

depending on whether the reference component is transferred from phase I into phase I1 or vice versa (see Fig. 1-58). YUthis the mole fraction of the upflow phase for a theoretical separation stage. The method used to determine stage efficiency is demonstrated in Fig. 1-58. The concentrations of the component of interest are experimentally determined in relation to the reference components of the entering and leaving fluxes Xu and xd.For a given

PhII

a)

b)

Fig. 1-57. Schematic of a countercurrent flow column with mass transfer from phase I1 to phase I (a) and an operating diagram to determine graphically the number of theoretical separation stages (b). EC Equilibrium curve BL Balance line n, n - 1,. . . Theoretical separation stages, refer to the loading of the key component in phases I and I1 as indices Y Loading of phase I1 with key component X Loading of phase I with key component

1.9 Steady-State Countercurrent Operation

85

current separation stage

Ph I Ph II a1

bl

Cl

Fig. 1-58. Definition of the stage or Murphree efficiency factor for a countercurrent flow column with the theoretical and practical maximum change in key component loading in a particular stage with mass transfer from phase I1 to phase I and vice versa.

EC Equilibrium curve BL Balance line Y,,, Phase equilibrium loading at X ,

equilibrium curve and balance line, the stage efficiency Egmis the ratio of the distance between these lines. The stage efficiency EI, is related to the downflow phase, and is thus similar to Eqs. (1-190) and (1-191). The stage efficiency varies from stage to stage according to the course of the equilibrium curve and balance line. It is dependent on the design of each separation stage, the process parameters in the column and the properties of the phases in contact. Usually the efficiency is determined experimentally. Empirical approximations may only be used to calculate the stage efficiency factor for a few simple cases. (Due to the nonideal mixture of the phases leaving the separation stage, the stage exchange rate Egmis the mean value of the local exchange rate Eg over the cross section of the stage, see Chapter 2.5.6.1. The latter value can only be calculated if the fluid mechanics of the phases at the cross section are known.) If E,, is the average of all stages, the approximate number of stages Np to be installed in t h e e m n is

Nt

Np = _ _

(1-192)

Egmm

where Nt is the number of theoretical stages, determined by computation or graphically. Where the slopes of the equilibrium curve and balance line are very different this approximation can lead to significant errors. For this case a “pseudo-equilibrium curve” between the equilibrium curve and balance line as shown in Fig. 1-59 may be introduced. Its course is determined by the vertical distances between the equilibrium curve and balance line divided by the stage exchange rate Egnl: a = Egm.b

(1-193)

The number of practical stages Np is determined by constructing steps between the balance line and pseudo-equilibrium curve. In Eq. (1-193), the exchange rate Egmis used repetitively in the interval X , + Xu. If trays are used as internals in a column, the theory of separation stages is particularly important for the determination of the actual column height (see Chapter 2.5.6.1).

1 Basic Concepts

PEC -4

5‘

/

Fig. 1-59. Determination of the number of practical separation stages. EC Equilibrium curve PEC Pseudo-equilibrium curve BL Balance line Y Loading of phase I1 X Loading of phase I

The values of n, and HETS depend on the form of the packing material, the type of packing (or the geometry in general), the operating conditions, and the properties of the two counterflow phases. nt can only be computed in particularly simple cases. The value is usually determined experimentally. If the number of theoretical stages Nt is determined, (for example by the graphical McCabe-Thiele method or by computation), and n, or HETS for the chosen packing are known, the height Z for mass and heat transfer of the packed column can now be calculated with Eqs. (1-194) or (1-195), respectively.

1.9.3 Calculation for Counterflow Columns

Graphical methods for the determination of the required number of theoretical stages discussed so far, are easy to use and are of practical importance for first approximations for separation processes of binary mixtures, or for systems with inert carrier mixtures and a single component to be transferred. If several components are transferred between the phases in contact, the number of separation stages can only be (1-194) calculated with the aid of a computer; graphical methods can only be partially apwhere n, is the number of theoretical plied. stages, which corresponds to the height Z of Methods of calculating the separation of a packed column and Nl is the number of multicomponent mixtures in counterflow theoretical stages that are effective along columns are based on the two “stage-tothe height of the packed column. Therefore, stage” methods by LEWISand MATHESON the reciprocal value HETP = HETS is the [1.99], or THIELE and GEDDES[1.100]. Both height of the packing, which has the same methods are iterative and usually exhibit usage as a theoretical stage (HETP: Height bad convergence, as well as needing a lot of Equivalent to one Theoretical Plate, HETS: computing time and power. Thus, until Height Equivalent to one Theoretical 1960, several approximation methods, or Stage). “short cuts” were developed and modified using simple assumptions to simplify the 1 HETP = (1-195) column design [1.101]. These short cuts are still of practical interest for an estimation nt

Heat and mass transfer in counterflow phases usually only takes place within the area of the trays. The theory of separation stages is also valid for packed columns. Since single stages cannot be distinguished, an evaluation number n, or its reciprocal value HETP or HETS is introduced

1.9 Steady-State Countercurrent Operation

87

been developed. However, a method giving sufficient and rapid convergence for a general design problem does not yet exist. Methods to calculate counterflow columns exactly are based on a column model. The column model gives a system of nonlinear equations, the task of each method is to solve this system. The equation system contains an overall mass balance and also a balance for each component, enthalpy balances, equilibrium relationships and the stoichiometric conditions for the sum of the concentrations for the theoretical stages of a counterflow column (see Fig. 1-60). In the equation system, the following nomenclature is used: Lj

Fig. 1-60. Model of a countercurrent column. i Flow rate of phase I (downflowing phase) SL Flow rate of side stream phase I F Feed 1, .. .j , . . .,n Theoretical separation stages (1 < j < n) j Stage index BA Balance area G Flow rate of phase I1 (upflowing phase) SG Flow rate of side stream phase I1 Q Heat flow

of the design of the column, and as first iteration variables for an exact check. If strong heat effects along the column height or pronounced nonideal behavior of the mixture occur, the short cuts fail. Since 1960, several methods for the exact calculation of a counterflow column have

flow rate of downflow phase leaving the jthstage Gj flow rate of upflow phase leaving the jthstage SLj flow rate of side stream leaving the jthstage (same state as the downflow phase) flow rate of side stream leaving SGj the jthstage (same state as the upflow phase) feed 6 .cj part of feed (same state as the downflow phase) FJ . (1 - c j ) part of feed (same state as the upflow phase) j index, characterizes the actual separation stage; starting at the top of the column, j is counted from 1 to n, from top to bottom heat flow from stage j Qj

6.

The column model in Fig. 1-60 is generally valid for rectification, counterflow absorption, and counterflow extraction. For rectification (see Chapter 2.5.2), the nomenclature used is : stage I (j=1)

condenser at the top of the column

88

1 Basic Concepts

SG, = O with total condensation SG,=O FI * E ] = O valid for normal cases, not F2=0 valid for a stripping column heat flow out of condenser Ql top product flux SL, vapor stream to condenser G reflux from condenser i, heat loss from an actual stage Qj (under approximately adiabatic operating conditions of the column) side stream, vapor SGj side stream, liquid SLj stage n

reboiler

F, =O

valid for normal cases, not valid for an enrichment column

(J =n>

F,+I= O heat supply to the reboiler bottom product flux vapor mixture generated in reboiler

- Qn

SL, Gn

and for counterflow absorption (see Chapter 3.5):

Fl . c1 F,+l

SG] SL,

feed of solvent needed to absorb the key component (washing liquid) (el = 1) feed of gas mixture to be treated ( E , + , = 0) flow rate of treated gas mixture leaving the column flow rate of loaded washing liquid leaving the column

Except for the listed flow rates, all other F,, SGj, SL,, and Qj are virtually zero. In an absorption column operated with an intermediate coolant as the washing liquid, Qj is substantially different than zero. For the case of counterflow liquid-liquid extraction (see Chapter 6.2.3.4)

F,.

E .,

F,+,

feed of heavy phase feed of light phase

SGl L,

s

light phase exit flow rate heavy phase exit flow rate

Usually no side streams are withdrawn ( s L j= 0, s G j = 0 ) and side feeds are not considered (5 = 0). At normal conditions, heat flux Qj through the sides or column walls is negligible. The following approximations are made for the calculation of a counterflow column : The concentration of the phases in one stage (on a tray) are uniform over the cross-sectional area of the column The phases leaving the stage are at phase equilibrium [the stage efficiency coefficient (enriching factor) may be used throughout the calculation] The column is operated at steady-state conditions (an exception being simulation of the start up of the column using dynamic methods; for example see [1.102]) In absorption and extraction processes, the assumption of phase equilibrium of the phases leaving the stage is less valid than in rectification processes. The exact calculation for a counterflow column using the theory of separation stages is generally carried out for rectification processes. The following discussion of a system of nonlinear equations, based on the column model, is particularly valid for rectification, but this may be applied to all other counterflow processes. Before the equation system can be formulated the variables used to describe the state of the counterflow system must be expressed. According to the column model in Fig. 1-60 they are

6

T.F

'.i

SLj s Gj

8;

i';

feed temperature of feed liquid fraction of the feed leaving side stream, liquid leaving side stream, vapor heat flux reflux flow rate

1.9 Steady-State Countercurrent Operation

Gj

T,, Pj x. 1 3 J.y y 1,J .. . z.[ , J.j L ’ z.1J.G

n

vapor stream temperature and pressure of the stages (on the column tray) molar fraction of key component i in the reflux stream L, and the vapor stream Gj molar fraction of the key component i in the liquid and vapor part of feed F j total number of separation stages (trays)

Equations showing the relationships of these variables are given in the following chapters.

1.9.3.1 Mass Balances The overall mass balance over the upper section of the column (see balance area BA in Fig. 1-60) gives J

(F, - SG, - SL,)

+

k=l

+ <+I

*

(1 - ~ j + l + ) Gj+l = Lj

89

If there are rn components in the mixture, rn individual mass balances must be calculated for each separation stage.

1.9.3.2 Phase Equilibrium Relationship For a theoretical separation stagej for each component i, the relationship between the phase equilibria of the leaving phases is

where KC is the equilibrium constant and is dependent on pressure, temperature and concentration. This relationship must be set up for m components and n stages.

(1-196)

where j is the stage number from 1 to n. A mass balance for each individual component i in each theoretical separation stage

1.9.3.3 Enthalpy Balances Enthalpy balances must be set up for each individual stage. For stage j ,

Fig. 1-61. Representation of separation stage j .

90

4

1 Basic Concepts

+ $+I * (1 - & j + I ) ~ , + I , G , F+ + L j p l 6;-1,t + G j + l h j + l , G *

Ej

*

hj,L,F

*

*

*

-(Gj+SGj)

*

hj,,-(Lj+SLj)

*

hj,t-Qj=O

(1-199)

5,,,

where hj+F and are the enthalpies of the liquid and vaporized part of the feed €$,and fij,L and fij,G are the enthalpies of the internal reflux Lj and vapor flow Gj, respectively (see Fig. 1-61).

1.9.3.4 Stoichiometric Conditions for the Sum of the Concentration at Each Equilibrium Stage

From the definition of the mole fraction for mixed phases it follows that rn

rn

Exlj = i

or

i= 1

or

m

m

i= 1

i= 1

C y i , j= 1

1.9.4 Kinetic Theory for the Counterflow Separation of a Mixture

i= 1

(1-200)

This sum is applied to each stage. The equation system is therefore defined by the mass balance, the enthalpy balance, the phase equilibrium relationships and the summation equation for the mole fractions. It contains n m mass balances, n enthalpy balances, n m phase equilibrium relationships and 2 . n summation conditions. Overall this gives n . (2m + 3) equations. Since the enthalpy hj and the equilibrium constant KCj are not linearly dependent on the state variables, particularly pressure and temperature, the equation system is nonlinear. In absorption columns, the number of unknown state variables equals the number of equations. In rectification columns three additional variables must be specified; ex-

-

amples are the reflux ratio at the top of the column and the mass flux of the top product. In extraction columns under normal operating conditions, the temperature profile is of little interest, and enthalpy balances need not be formulated. The solution to the equation system can be obtained using an iterative procedure. An overview and comparison of different iterative procedures are given in [1.102- 1.1041. The solution of the equation system gives the profile of the molar fluxes, the concentration profile and the temperature profile of a counterflow column. Finally the dimensions of the counterflow column need to be determined and the column must be integrated into the separation plant.

If column internals, such as filling or packing material, are used to increase the phase boundary area, then the phases are in constant contact over the column height; this is also true for spray columns. The constant contact of the phases should therefore be considered in column calculations. Instead of the separation stage theory, where the phase contact only takes place within the stages, the kinetic theory for the counterflow separation of a mixture is used. This theory allows the calculation of the actual height of the column where heat and mass transfer occur. The procedure is continuous, does not take single stages into account, and is based on the mass transfer relationships discussed in Chapter 1.7.3. This procedure is also used when the stage efficiency is very small, for example in physical absorption processes.

1.9 Steady-State Countercurrent Operation

1.9.4.1 mo-Directional Mass Transfer Between Phases

Figure 1-62 shows a section of a continuously operated column with two counterflow phases in constant contact. Mass is transferred between both phases, in either direction (for example with rectification). The flow rates of the phases G and L are constant along the column height if the molar vaporization enthalpy of the components is equal, and no side streams appear, no heat losses need to be considered and the mixing enthalpy is negligible. This is approximately true for rectification processes where enriching and stripping sections are treated separately (see Chapter 2.5.2) and no side streams occur. If there are side streams, the column must be calculated in a section wise manner. It follows that for the flow rate dri of the key component in the height element dz, the mass balance is dri = 6 . dy = L . dx

91

tration scale, and the mass transfer coefficient k, is related to the upflow phase. dA is the contact area in height element dz for mass and heat transfer. The specific volumetric surface a of the column packing is the ratio of the surface of the packing to the volume in m2/m3 and AQ is the cross-sectional area of the column. Hence,

dA = q . a . A Q - d z

(1-203)

(Generally, the specific surface a is not completely available for heat and mass transfer; it has to be corrected with an efficiency factor q < 1. q is the ratio of the physical wetted surface a, and the theoretical wetted surface a. q is dependent on the type or method of packing, the properties of the phases in contact, the distribution of the phases across the cross-sectional area of the column and the loading of the cross section by the phases.) From Eqs. (1-201)-(1-203) it follows that

(1-201)

Combining this with Eq. (1-174) dri gives dli=k,*(y-y*).dA

(1-202)

where the mole fraction y is used (instead of the molar concentration c) as the concen-

G Y+

-'

L

X+

d!-. - -~- _ _dx

-.

"+-

G Y

L X

Fig. 1-62. Height Z of a countercurrent column for mass and heat transfer.

Integrating with respect to dz gives the height Z of the column for mass and heat transfer.

If G and k, are not constant along Z , this must be taken into account during the integration. dy is the change of concentration in the upflow phase, in the height element dz, effected by the driving force y - y*. The integral of Eq. (1-205) is, according to CHILTON and COLBURN [1.105] the Number of Transfer Units (NTLr). NTU,, =

dy

Yn ~

yw

Y -Y*

(1-206)

92

1 Basic Concepts

NTU,, gives a number of transfer units describing how often the driving force y - y * is needed in the overall interval ya - y w , and the concentration differences between the entry and exit (z = 0 and z = 2) of the column. The overall mass transfer resistance is considered which is expressed by the index og (overall gas) and is related to the upflowing phase. The Height of a Transfer Unit,HTU,, is given by HTU,,=

G

kg*

Y,l*

a . AQ

= HTU,,

. NTU,,

(1-210)

Z

L

k,*q*a.AQ

= HTU,,.

NTU,,

Z=

LT

dX

(1-214)

(1-208)

(1-209)

=

(1-213)

and

Analogously, with respect to the downflow phase (index 01)

HTU,I

dY

Z=

(1-207)

The height of the column for heat and mass transfer is thus Z

of the key component in the inert part of the fluxes X and Y can thus be used as a concentration scale for the determination of the column height Z . With small concentrations, X and Y are almost equal to the molar fractions x and y. Applying this to the derivative calculated in the previous chapter, gives

(1-211)

At higher concentrations of the key component, the molar fraction may be expressed as

Introducing this expression, Eq. (1-214) becomes

dx

Z = q * LT . -~)~.(,y*-x) a m A Q*x,T k l (1 (1-212)

1.9.4.2 One-Directional Mass lkansfer

In absorption, adsorption, and extraction processes, substances are usually only transported from one phase which carries the key component, into another phase which takes the component. The phase fluxes G and L are not constant along the column height 2, although the inert parts of the phase fluxes G, and L, are constant. The concentration

(1-216)

which may also be applied when the mass transfer coefficient k , is dependent on the concentration. If the balance line and equilibrium curve are parallel, NTU,, and NTU,, are identical. NTU,/ and NTU,, are then equal to the number of theoretical stages determined using a graphical method. It is sufficient with an approximately linear equilibrium curve and balance line to replace the driving force y - y * or x * - x in

1.9 Steady-State Countercurrent Operation bl

a)

t

93

ya

1.

Yl

Fig. 1-63. Graphical determination of the number of transfer units NTU,,. Area under the curve A 1 NTU,, (Eq. 1-213). a) Determination of the driving force Y - Y* b) Graphical evaluation of the integral NTU,, BL Balance line

EC Equilibrium curve

the column. Methods of calculating HTU are given in the discussion of the individual separation processes. Direct experimental determination of k . a, or fi a, under con(Y - Y*), - (Y - r*), (1-217) ditions as close as possible to the operating conditions is often required. (Y - Y*), (The values of the height of a transfer In (Y - Y *), unit HTU,, and HTU,, may be calculated from the mass transfer coefficients fig and and it follows that PI related to the heights HTU, and HTU,. This is analogous to the calculation of the yu dy overall mass transfer coefficient k, based " JO'(1-218) NTU,,= on the mass transfer coefficients P, and PI, y , y - Y* (Y -Y*)/m for example, the height of a transfer unit HTU,, is given by Thus a transfer unit exists if the change of concentration Ay equals the mean driving GT HTU,, = HTU, + m T- HTU, (1-219) force causing the change. LT The integrals NTU,, and NTU,, can simply be evaluated numerically or graphi- with cally, as shown in Fig. 1-63. Graphical GT methods approximating NTU,, and NTU,,, HTU, = (1-220) Pg'r'a'AQ respectively were introduced by BAKER [I .106]. L T The heights of the transfer units HTU,, (1-221) HTU, = and HTU,, largely depend on the mass PI * r . a * A Q transfer coefficients k, and k,, the specific exchange area a and the molar fluxes of the where rn is the slope of the equilibrium phases through the cross-sectional area of curve.) the integral, by the logarithmic mean, ( y - y *),m or (x* - x),~,at both ends of the column. For example for the gas phase,

S --

-

-

-

94

1 Basic Concepts

1.10 Steady-State Crossflow

Operation Figure 1-64 shows a cascade of three stages operated at steady-state in a crosscurrent flow process where heat and mass transfer occur.

ration processes but rarely in solid-liquid and liquid-liquid extraction, drying processes, or adsorption.

t

I

L

Yl v2 Y3

x Fig. 1-64. Steady-state cross-flow operation of 3 stages.

Phase I flows through each stage in turn. Phase I1 is fresh feed to each separation stage. Within the stages, both phases are in cocurrent flow, with mass transfer from phase I into phase 11. A mass balance over the first stage of the crossflow cascade, according to Eq. (1-178) and the notation in Fig. 1-64, gives i T

*

(xa- X I ) = G,

1

*

( 5 - Ya)

(1-222)

For the case of constant inert or carrier fluxes, this equation gives a straight line when plotted on an X, Y graph, of slope - L T / G T , ; this is also true for the other stages. Figure 1-65 shows the operating diagram for the crossflow cascade, including the equilibrium curve and balance line. (The of the balance lines are . slopes . - L T / G T l , - L T / G T z , and -LT/G;7;.3). The crossflow of phases in contact, as shown in Fig. 1-64, is used in thermal sepa-

4

Fig. 1-65. Loading or operating diagram for a steady-state cross-flow cascade consisting of three separation stages. BL,, BL,, BL, Balance lines for stages 1, 2, 3 EC Equilibrium curve Loading of phase I1 by the key Y component Loading of phase I by the key X component

1.11 General Procedure to Design Equipment for the Thermal Separation of Mixtures The design of equipment for the thermal separation of mixtures and its integration into a whole separation unit for a chemical plant follows a certain procedure, shown in Table 1-19. In this book, the chemical engineering aspect of the design of separation equipment is discussed. Literature for stress calculations for parts of the plant may be found in the field of technical mechanics and plant construction [1.107].

1.11 General Procedure to Design Equipment for the Thermal Separation

95

Table 1-19. Planning concepts for thermal separation process units. *

heat transfer by graphical or short-cut methods, Iterative calculations, Correction using Murphree stage efficency coefficients

apparatus length for mass and heat transfer with the help of kinetic

* The order of individual steps in the table does not necessarily mean a chronological order. Generally, according to a time schedule (critical path analysis, Gantt plan) steps are carried out in parallel as much as possible.

96

1 Basic Concepts

References

[ 1.181 SCHILLING, R. : VGB Kraft werkstech. 60

(1980) 9, 695-705. [l. 191 SATTLER,K. : Thermische Trennverfahren. Aufgaben und Losungen. Ausle[1.11 SCHULZE,J., and HASSAN,A.: Methoden gungsbeispiele. Vogel-Verlag, Wiirzburg der Material- und Energiebi~anzierungen 1979. bei der Projektierung von Chemieanla- [ 1.201 Physiochemical fundamentals. gen. Verlag Chemie, Weinheim 1981. [1.21] BENDER,E., BLOCK,U.: VerfahrenstechA.: Grundlai1.21 BENEDEK,P., and L,ASZU~, nik 9 (1975) 3, 106-111. gen des Chemieingenieurwesens. VEB [ 1.221 HLATVA,K. : Collect. Czech. Chem. ComDeutscher Verlag fur Grundstoffindumun. 37 (1972), 4005-4007. strie, Leipzig 1965. [ 1.231 FRANCIS,A. W. : Liquid-liquid-Equili.31 HASSAN,A., Dissertation, TU Berlin briums. Interscience Publishers J. 1979. Wiley&Sons, New York 1967. v.41 KOGL,B., and MOSER,F. : Grundlagen der [1.24] SEIDELL, A., and LINKE,W. F.: Solubility Wrfahrenstechnik. Springer Verlag, Wien of Inorganic and Organic Compounds, 2 1981. Vols. Van Nostrand, Princeton 1958 and [1.51 MULLER,W. H., and MUSCHIK,W.: J 1965. Non Equil. Thermodyn. 8 (1983) 1, [1.25] D'ANs, J., and LAX, E. (eds.): Taschen29-66. buch fur Chemiker und Physiker, Vol. 1. P. : Verfahrenstechnik 13 U.61 GRASSMANN, Springer-Verlag, Berlin 1967. (1979) 1, 28-31. [1.26] STEPHEN,J., and STEPHEN,T. (eds.): So~1.71ARNOLDand BARTMANN:Chem. Ing. lubilities of Inorganic and Organic ComTech. 53 (1981) 7, 497-507. pounds, 2 Vols. Pergamon Press, Oxford W. : Energetische Analyse [ I 3 1 FRATZSCHER, 1964. von Stoffubertragungsprozessen. VEB [I .27] LANDOLT-BORNSTEIN, Vol. 11, part 2b, 2c. Deutscher Verlag fur GrundstoffinduSpringer Verlag, Berlin from 1950. strie, Leipzig 1980. [ 1.281 FRANCIS, A. W. : Handbook for Compo[I .91 SCHAFER,W. : Verfahrenstechnik 4 (1970) nents in Solvent Extraction. Gordon and 8, 352-358. Breach Science Publishers, New York [1.10] WOZNY,G., FETT, F., and CREMER,H.: 1972. Verfahrenstechnik 17 (1983) 6, 375-381 D. M., BRADY,B. L., and [1.29] HIMMELBLAU, and 7, 433-439. MCKETTA,J. J. : Survey of Solubility DiaA. W., HUTCHINSON, H. P. [1.11] WESTERBERG, grams for Ternary and Quaternary LiMOTARD,R. L., and WINTER,P.: Process quid Systems. Bureau of Engineering ReFlowsheeting. Cambridge University search, University of Texas 1959, Special Press 1979. Publication No. 30. T. Verfahrens- [ 1.301 SORENSEN, [1.12] BMAR,L., and PILHOFER, J. M. and ARLT,W. : ,,Liquidtechnik 7 (1973) 2, 53-57. Liquid-Equilibrium Data Collection." [1.13] KAUFMANN, F., HOFFMANN, U., and HOFIn: DECHEMA Chemistry Data Series, 3 MANN,H.: Chem. Ing. Tech. 45 (1973) 7, Vols. Frankfurt from 1979. 450-455. [ 1.311 GRASSMANN, P. : Physikalische Grundla[1.14] FUTTERER,E., and PATTAS,E.: Chem. gen der Verfahrenstechnik. Verlag Salle Ztg. 98 (1974) 9, 438-445. und Sauerlander, FrankfurUMain, Aarau [1.15] PIERUCCI,S. G., BIARDI,G., RANZI,E., 1983. and DENTE,M.: Chim. Ind. Milan 62 [I .32] BEST,R., SPINGLER, E.: Chem. Zng. Tech. 44 (1972) 21, 1222- 1226. (1980) 3, 193-199. [1.16] KETCHUM,R. G.: Chem. Ing. Tech?.44 [1.33] FREUNDLICH, H. : Coffoidand Capillary Chemistry. London 1926. (1972) 7, 457-462. [1.17] LINNHOFF,B., and RJRNER,J.: Chem. [1.34] LANGMUIR,J.: J. Am. Chem. SOC. 38 (1916) 221. Ind. Duesseldorf 104 (1981) 9, 544-550.

References [1.35] BRUNAUER, S., EMMET,H. H., and TELLER, E. : J. Am. Chem. SOC. 60 (1938) 309. [1.36] BRUNAUER, S., DEMING,L. S., DEMING, W. G., and TELLER,E.: .lAm. Chem. SOC. 62 (1940) 1726. [1.37] DUBININ,M. M. : Chem. Rev. 60 (1960) 235-241. [1.38] JUNEEN, H.: Staub Reinhalt. Luft 36 (1976) 7, 281-287. [1.39] MERSMANN, A.: Thermische Erfahrenstechnik. Springer-Verlag, Berlin, Heidelberg 1980. [ 1.401 MERSMANN, A., MUNSTERMANN, U., and SCHADL, J. : Chem. Zng. Tech. 55 (1983) 6, 446-458. [1.41] HOPPE, H. and WORCH,E.: Wiss. Z. Tech. Hochsch. Chem. Carl Schorlemer Leuna Merseburg 23 (1981) 3/4, 418-428. [1.42] MYERS,A. L., and PRAUSNITZ, J. M. : AIChE J. 11 (1965) 1, 121-127. [1.43] COSTA,E., SOTELO,J., CALLEJA, G., and MARRON, C. : AZChE J. 27 (1981) 1, 5- 12. [1.44] KAST,W., and DREHER,H. : Chem. Zng. Tech. 51 (1979) 12, 1245. [1.45] HOPPE,H., and WORCH,E.: Chem. Tech. Leipzig 31 (1979) 9, 464-467. [1.46] SEIDELL, A., and LINKE,W. F. : Solubility of Inorganic and Organic Compounds,2 Vols. Van Nostrand, Princeton 1958 and 1965. [1.47] D'ANs, J., and LAX,E. (Hrsg.): Taschenbuch fur Chemiker und Physiker, Vol. 1. Springer, Berlin 1967. [1.48] STEPHEN H., and STEPHEN,T. (eds): Solubilities of Inorganic and Organic Compounds, 2 Vols. Pergamon Press, Oxford 1964. [1.49] LANDOLT-BORNSTEIN, Vol. 11, 2b, 2c. Springer, Berlin, from 1950. [ 1.501 International Critical Tables. McGrawHill Book Comp., from 1933. [1.51] PRAUSNITZ,J. M., and SHAIR,F. H. : AZChE. J. 7 (1961), 862. [1.52] FRIEND,L., and ADLER,S. B.: Chem. Eng. Prog. 53 (1957) 452. [1.53] BATTINO, R., and CLEVER, H. L., Chem. Rev. 66 (1966) 395. [ 1.541 SEIDEL,A. : Solubilities of Inorganic and Metal-Organic Compounds. Van Nostrand, New York 1958. [1.55] WILHELM, E., and BATTINO,R.: Chem. Rev. 73 (1973) 1.

97

[1.56] WILHELM, E., BATTINO, R., and WILCOCK, R. J.: Chem. Rev. 77 (1977) 219. [1.57] NAKAHARA,T., and HIRATA, M.: J. Chem. Eng. Jpn. 2 (1969) 2, 137-142. [1.58] HALA,E., PICK,J., FRIED,V., and VILIM. 0.: Gieichgewicht Fliissigkeit - DampJ Akademie-Verlag, Berlin 1960. [1.59] HALA,E., PICK,J., FRIED,V., and VILIM, 0.: Vapour-Liquid-Equilibrium.Pergamon Press, Oxford 1967. [1.60] COTTRELL, F, G. : J. A m . Chem. SOC.42 (1919) 721-729. [1.61] Publications of Fischer Labor- und Verfahrenstechnik, Meckenheim. [1.62] NULL,H. R.: Phase Equilibrium in Process Design. Wiley-Interscience J. Wiley and Sons, New York 1970. [ 1.631 KING,M. B. : Phase Equilibrium in Mixtures. Pergamon Press, Oxford 1969. [1.64] PRAUSNITZ, J. M.: Molecular Thermodynamics of Fluid-Phase-Equilibria. Prentice Hall Inc., Englewood Cliffs N.J. 1969. [1.65] PRAUSNITZ, J. M., and ECKERT,C. A. : Computer Calculations for Multicomponent Vapor-Liquid-Equilibria. Prentice Hall Inc. Englewood Cliffs N.J. 1967. [1.66] PRAUSNITZ, J. M., and CHUEH,P. L. : Computer Calculations for High-Pressure Vapor-Liquid-Equilibria. Prentice Hall Inc., Englewood Cliffs N.J. 1968. [ 1.671 CHU, J. C. : Distillation Equilibrium Data. Reinhold Publishing Corp., New York 1950. [ 1.681 CHU, J. C. : Vapor-Liquid Equilibrium Data. J. W. Edwards Publishers Inc. Ann Arbor Mich. 1956. [1.69] KOGAN,V. B., and FRIEDMA", V. M.: Handbuch der DampfFliissigkeitsGleichgewichte. VEB Deutscher Verlag der Wissenschaften, Berlin 1961. [1.70] PRAUSNITZ,J. M., and GMEHLING,J.: Thermische Erfahrenstechnik ' Phasengleichgewichte.Krausskopf-Verlag, Mainz 1980. [1.71] STEPHAN, K.: Chem. Zng. Tech. 52 (1980) 3, 209-218. [1.72] GMEHLING, J., and ONKEN,U.: ,,VaporLiquid Equilibrium Data Collection. " In: DECHEMA-Chemistry Data Series, from 1977.

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1 Basic Concepts

[I .86] ROCK, H. : Destillation im Laborato[1.73] KOGAN,V. B., FRIEDMAN, V. M., and KArium. Extraktive und azeotrope DestillaEAROV, V. V. : DampjWussigkeifs-Gleichtion. Steinkopff, Darmstadt 1960. gewichte. Verlag Nauka, Moskau 1966. [1.87] HAASE,R. : Thermodynamik der MischI., POLAK,J., and [1.74] HALA,F., WICHTERLE, phasen. Springer, Berlin 1956. BOUBLIK,T. : Vapor-Liquid Equilibrium [1.88] ROWLINSON, J. S.: Nature213 (1967) 440. Data at Normal Pressures. Pergamon [1.89] BARROW, G. M. : Physikalische Chemie, Press, Oxford 1968. part 111: Mischphasenthermodynamik, [ 1.751 Kux, C. : ,,Dampfdriicke von Mischsyste(eds.): Elektrochemie, Reaktionskinetik. Bohrmen. '' In: LANDOLT-BORNSTEIN, mann, Vieweg, Wien 1971. 6th Ed. Vol. 11. part 2a, pp. 336- 767. [1.90] ECKERT,E. R. G., and DRAKER,R. M.: Springer Verlag, Berlin 1960. Analysis of Heat and Mass Transfer. [1.76] STAGE,H., and FALDIX,P.: Fortschritte McGraw-Hill Book Comp. New York der Verfahrenstechnik. Verlag Chemie, 1972. Weinheim, from 1954. [1.91] FULLER,E. N., SCHETTLER, P. D., and [1.77] DELLRICH,I. R., P ~ C K E RU., J., and GIDDINGS,J. C.: Ind. Eng. Chem. 58 KNAPP,H. : Vapor-Liquid Equilibria. A (1966) 5, 19-28. Bibliography of Published Data of Multi[1.92] JOST,W.: Diffusion in Solids, Liquids, component Systems Containing CompoGases, Academic Press, New York 1960. nents with Normal Boiling Points Lower [1.93] HIGHBIE,R.: Trans. Inst. Chem. Eng. 31 than 350 K. (Bibliographie der Veroffent(1935) 365. lichungen seit 1900 bis 1972. Nachweise von 350 Binar-, 90 Ternar- und 30 Mehr[1.94] DANKWERTS, P. V. : Znd. Eng. Chem. 43 (1951) 1460. komponentensystemen). Institut fur [I .95] BRAUER,H. : Stoffaustausch einschlieJThermodynamik. TU Berlin 1973. lich chemischer Reaktionen. Sauerlan[1.78] Dechema material data office. der, Aarau 1971. [ 1.791 RMMERMANS, J. : The Physico-Chemical 11.961 MERSMANN, A. : Stoffubertragung. SprinConstants of Binary Systems in Concenger-Verlag Berlin, Heidelberg 1986. trated Solutions. Interscience Publishers, [ 1.971 SCHLUNDER, E.-U. : Einfuhrung in die New York/London 1960. Stoffubertragung. Thieme-Verlag, Stutt[1.80] CHEMICAL ENGINEERING SOCIETY (JAPAN) gart 1984. (eds.) : Physikalische Konstanten. Maru[1.98] MCCABE,W. L. and ~ I E L E E. , W.: Ind. Zen-Verlag, Japan, from 1963. Eng. Chem. 17 (1925) 605. [1.81] STAGE,H.: Bibliographie von ca. 2000 [I .99] LEWIS,W. K., and MATHESON, G. L.: Binursystemen. Destillationstechnik Dr. Ind. Eng. Chem. 24 (1932) 494. H. Stage, Koln-Niehl. R. L.: Ind. (1.821 HORSLEY, I. H.: ,,Aceotropic Data." In: [I .I001 ~ I E L EE., W., and GEDDES, Eng. Chem. 25 (1933) 289. Advances in Chemistry; Series Vol. 6 and 35. American Chemical Society. Wash- [1.101] FENSKE,M. R.: Ind. Eng. Chem. 24 (1932) 482. UNDERWOOD, A. J. V.: ington 1952 and 1962. Chem. Eng. Prog. 44 (1948) 8, 603-614. [1.83] OGORODNIKOV, S. K., LESTEVA, T. M., GILLILAND, E. R.:Ind. Eng. Chem. 32 KOGAN,V, B.: Azeotrope Gemische, Ver(1940) 1220. lag Chemie, Leningrad 1971. [1.84] BUSCH,A.: ,,Binare Systeme - azeo- [l. 1021 KETCHUM,R. G. : Chem. Ing. Tech. 43 (1971) 5, 264-269. trope Gemische." In: Landolt-BornEDULJEE,H . E.: Hydrocarbon Process. stein (eds.): 6th Ed. Vol. 2, part 2a, 54 (1975) 9, 120-122. pp. 663/711. Springer-Verlag, Berlin ERBAR,R. C. JOYNER,R. S., and MAD1960. DOX, R. N.: PetroKhem Eng. 33 (1961) 11.851 WICHTERLE, I., LINEK,J., and HALA,E. : 3, C 19-C 22. Vapor-Liquid Equilibrium Data BiblioWINN, F. W.: Pet. Refiner 40 (1961) 4, graphy. Elsevier Scientific Publishing ~ . . Comp., Amsterdam 1973. 153- 155.

References [1.103] GELBE, H., and NOMINE,H.: Verfahrenstechnik 5 (1971) 10, 429-435. [1.104] NEUMANN, K. K.: Erdoel, Kohle, Erdgas, Petrochem. Brennst. Chem. 26 (1973) 4, 198-202. [I .lo51 CHILTON,R. H., and COLBURN, A. P. : Ind. Eng. Chem. 27 (1935) 255. [1.106] BAKER,T. C.: Ind. Eng. Chem. 27 (1935) 977. [I. 1071 KLAPP,E. : Apparate und Anlagentechnik. Springer-Verlag, Berlin, Heidelberg 1980. SCHWAIGERER,S. ; Festigkeitsberechnung im Dampfkessel-, Behalter- und Rohrleitungsbau. Springer-Verlag, Berlin 1979. BTZE, H. : Elemente des Apparatebaus. Springer, Berlin 1967. AD-Merkblattter der Arbeitsgemeinschaft Druckbehaltec VdTUV, Essen. RABALD,E., and BRETSCHNEIDER, H.: DECHEMA-Werkstofftabelle.Verlag Chemie, Weinheim, from 1954. PIATTI,L. : Werkstoffeder chemischen Technik. Sauerlander, Aarau.

99

WAGNER,W. : Apparate- und Rohrleitungsbau. Vogel-Buchverlag, Wiirzburg 1984. [ 1.1081 JORDAN, D. G. : Chemical Process Development, 2 Vols. Interscience Publishers, J. Wiley & Sons, New York 1968. [1.109] MACH, E.: Planung und Einrichtung chemischer Fabriken. Sauerlander, Aarau 1972. BERNECKER, G. : Planung und Bau verfahrenstechnischer Anlagen. VDI-Verlag, Diisseldorf 1984. ULLRICH, H. : Anlagenbau. Georg Thieme-Verlag, Stuttgart 1983. [ 1.1101 LUDWIG,E. E. : Applied Process Design for Chemical and Petrochemical Plants, 3 Vols. Gulf Publishing Comp., Houston 1965. [1.111] WEHDE, K.-H., and STICHLMAIR,J.: Chem. Zng. Tech. 57 (1985) 4, 348-349. H.: Chem. Ing. Tech. 55 [1.112] HORMEYER, (1983) 1, 54-55. [ 1.1131 WETZLER,H. : Kennzahlen der Wrfahrenstechnik. Hiithig-Verlag, Heidelberg 1985.

Thermal Separation Processes: Principles and Design Klaus Sattler, Hans Jacob Feindt copyright 0VCH

Verlagsgesellschaft m h H , 1995

2 Distillation and Partial Condensation

2.1 Concepts of Simple Distillation, Rectification and Partial Condensation Distillation is a process for the thermal separation of liquid mixtures. The process may be used when the boiling points of the components in the mixture are significantly different, provided the substances are thermally stable under the operating conditions. The following is typical for a distillation process: by supplying heat, approximately isobarically, some of the liquid mixture is vaporized. The vapor is then distilled off and condensed after separation from the remaining liquid mixture, forming a second phase, the vapor phase. The low-boiling substances of the mixture are enriched in the vapor phase (case: relative volatility ai,j > 1). The “selective” vapor phase generally has a different composition to the liquid phase (an exception being an azeotropic mixture). By separating the vapor phase from the remaining liquid mixture a partial separation is carried out and since the vapor phase is enriched with the more volatile components, the less volatile components are concentrated in the distillation residue. In the process of partial condensation some of the vapor mixture is condensed by the removal of heat. Low-boiling components are therefore enriched in the remaining vapor, while the condensate of the additional “selective” phase contains the

higher-boiling components of the original vapor mixture. Figure 2-1 shows the effects of partial distillation and partial condensation on a binary mixture in a bubble point temperature diagram. Heating this mixture, which has a mole fraction x,, of the lower-boiling component and a temperature t9,, to the temperature dBC gives a liquid and a vapor phase which are in equilibrium. B is the state point of the remaining liquid phase, which is reduced in the lower-boiling component (xB < x,). The distillate (point C, yc > x ,) is enriched with lower-boiling component. At the point where boiling begins, (state point S, boiling point temperature rSS) the change of state follows the bubble point curve from S to B. Cooling the

xB

xA

x.y

-

YC xE

YO

YF

Fig. 2-1. Partial distillation and condensation in a bubble point temperature diagram. BC Boiling point curve DC Dew point curve

102

2 Distillation and Partial Condensation

vapor mixture, by means of partial condensation and at phase equilibrium, (from dD to L9E,F) gives a partial condensate containing less lower-boiling component than the original vapor (point El x, < yo), while the remaining vapor (point F, y, > yo) is enriched with the lower-boiling component. This change of state follows the dew point curve from D to F. In contrast to the evaporation of solutions, where only the solvent vaporizes and the partial pressure of the soluted substance is negligible, in distillation processes all components found in the liquid phase are also present in the vapor phase. The fraction of each component in the vapor phase depends upon the “effort” required to escape from the liquid into the vapor phase. The relative volatility is a measure of this “effort”, the separability of a mixture by distillation (discussed in Chapter 1.5). Since the relative volatility increases with decreasing pressure, the separation efficiency of a distillation process is increased by decreasing the operating pressure. Distillation processes may be operated continuously or batchwise. The operating pressure may be low (10’ Pa-100 Pa), medium (10 Pa-0.1 Pa) or even high vacuum (0.1 Pa- l o p 3 Pa) (molecular distillation) to avoid thermal damage to the mixture, or to increase efficiency. Special types of distillation are equilibrium or flash distillation (flash evaporation) and carrier distillation processes. A simple single distillation stage gives an imperfect separation of liquid mixtures. The single distillation stage is repeated in a counterflow distillation process, or rectification, to give separate components of the desired purity. In rectification processes the generated vapor is guided through the rectifying column in counterflow to parts of its condensate (Fig. 2-2). If the relative volatility a > 1, at the intensive contact of the counterflow phases

.I

ss

Fig. 2-2. Continuous countercurrent distillation unit. RC Rectifying column ES Enriching section, rectifying column SS Stripping section, exhausting column C Condenser DC Distillate cooler DR Distillate receiver R Reboiler RC Residue cooler RR Residue receiver F Feed Sidestream G Vapor to condenser Overhead product, light ends R Reflux A Bottom product, residue, heavy ends LC Lower boiling component H C Higher boiling component

s

2.2 Discontinuously and Continuously Operated Simple Distillation

(vapor and reflux) heat and mass transfer occurs. In binary processes, the less volatile component of the vapor condenses into the liquid phase. In adiabatic operation of the column, the heat released during condensation is used to vaporize the more volatile components from the reflux. Thus, the vapor flow up the column is enriched with the more volatile components, whilst the reflux is enriched with the less volatile components from top to bottom. For example, when a nonazeotropic binary mixture is rectified, the vapor leaving the top of the column will be nearly pure lower-boiling component, whereas at the bottom of the column, the liquid leaving will be nearly pure heavy boiling component. Heat and mass transfer can be intensified inside the column by using internal elements, such as trays and packing. These internals increase the contact time between the phases and the interfacial area. A separation by rectification of a multicomponent mixture cannot be completely carried out in a single column. Several columns connected in series are required to give a “pure” product for each component. For example, for m components without an azeotropic point, m-1 columns are required to give complete separation. Rectification processes may be operated continuously and discontinuously. Under adiabatic conditions the process can be operated at normal pressure, underpressure, and overpressure. Azeotropic mixtures are treated using azeotropic or extractive rectification. For special cases nonadiabatic, thermal rectification is used. The operation conditions and the type of internals used in the rectification column depend on the behavior of the mixture during separation and the properties of the components present.

103

2.2 Discontinuously and Continuously Operated Simple Distillation, Flash Distillation 2.2.1 Discontinuously Operated Simple Distillation In discontinuous simple open distillation (discontinuous partial distillation, Rayleigh Distillation) the distillation still is charged with a liquid mixture. Slow heating to the boiling point partially vaporizes the liquid. The vapor becomes enriched with the more volatile components and is withdrawn from the distillation still. After condensing in a condenser the distillate is stored in a distillate receiver (Fig. 2-3).

DR

Y Fig. 2-3. Batch distillation unit.

DS Distillation still Distillate condenser cooler DR Distillate receivers C

If the distillate is collected in several different distillate receivers, over a time period, the process is known as fractional simple distillation. The distillate is subdivided into several different batches, called cuts, of different purities. The fraction of the less volatile components in the batches

104

2 Distillation and Partial Condensation

increases with time, whilst the fraction of the more volatile component remaining reduces. Fractional simple distillation is therefore only used for partial separation of multicomponent mixtures. During discontinuous simple distillation, the concentration in the still, the distillate concentration and the temperature all change as the process proceeds. Figure 2-4 shows these changes as a function of time.

t

d

Fig. 2-4. Concentrations and temperatures at discontinuous distillation as a function of distillation time f. xn Still concentration x, Distillate concentration 8, Still temperature d D Distillate temperature

Masses and mass fractions may be used in place of moles and mole fractions. It is thus only possible to state xQm if B,, xBa, and xBw and the distillation residue B, are known. A differential mathematical approach must be used to obtain B,, this is discussed as follows. During the distillation, at any time, the molar content of the distillation still is B with mole fraction x of the key component. After distilling a partial amount d B with mole fraction y of the key component, the residue in the still becomes B - d B with mole fraction x - dx. Neglecting the product dB . dx and separating the variables x and y in the mass balance, a differential equation for the discontinuous simple distillation is derived. This equation was first introduced by LQRDRAYLEICH and is called the Rayleigh equation dB - dx B y-x Integration between the limits B,, B,, on the left-hand side of Eq. (2-2) and with the limits x,,, x, on the right-hand side (2-3)

The integral in Eq. (2-3) can only be evaluated if the vapor concentration y ( x ) is The liquid mixture being separated B, known as a function of the concentration x, with a molar fraction x, of the more volain the still during the distillation period. tile component (key component), is distilled Numerical results from experiments are during a period of time to obtain a distillagenerally used. tion residue B , with x , ~ . In the distillate Assuming phase equilibrium between the receiver the number of mols of distillate is liquid and the vapor in the still, the vapor D = B, - B, with a mean fraction x ~of , concentration ~ y ( x ) is described by the equithe key component. From an overall matelibrium curve. The integral in Eq. (2-3) can rial balance of the distillation unit the mean now be evaluated using numerical or graphdistillation concentration is ical means (Fig. 2-5). The assumption of phase equilibrium between vapor and liquid in the still is improved when the contents of the still are intensively agitated, and when

2.2 Discontinuously and Continuously Operated Simple Distillation

:1-

I

k

1 Yl-xl

xBw

x1

105

Fig. 2-5. Determination of the definite integral in Eq. (2-3). a) Equilibrium diagram b) Graphical integration method EC Equilibrium curve y Molar fraction of the key component in the vapor x Molar fraction of the key component in the liquid

XBa

a)

less distillate compared with the contents of the still, is distilled during a period of time. A special case occurs when the relative volatility 01 is constant in the concentration range of interest, xBa> x > xBw. Provided that there is phase equilibrium between the contents of the still and the distillate vapor, the amount of distillate D can be expressed as

-.

(2-4)

For a binary mixture the mean distillate concentration of the lower-boiling component is then

(2-5)

In order to assess the volume of the distillation still the following time periods must be considered: the time required to charge and

discharge the still, the time required to heat the still contents to the boiling point, the distillation time, the time to cool the distillate residue, and if necessary, the time required to form the vacuum and to bring the apparatus back to atmospheric pressure. If the distillation unit is part of a continuous chemical plant, an additional buffer is also required. The minimum size of the distillation still and buffer depends on the liquid volume fed from the previous units of the plant during the total time period required to operate the distillation unit. Free vapor space must be added to the fixed volume of the distillation still and an extra safety margin to the buffer volume. Therefore, the minimum volume V of the still is

where

v, tg

VD

volumetric flow rate of the feed mixture at its boiling (m3/h) overall batch time volume of the free vapor space

Distillation and Partial Condensation

molar mass and density (with

where

diameter and height of the free vapor space of the still

C

xB, at L9,) of the fed mixture

eg, e, density of the distillate vapor and

The diameter d of the distillation still is then calculated from the relationship d=

-1 ~

n: a3600

eg*w

(2-7)

Z , is the minimum height of the vapor space (see Fig. 2-3)

where diameter of the distillation still maximum flow of distillate vapor during the distillation period (kmol/h) Mg molar mass of the distillate vapor eg density of the distillate vapor w maximum rising velocity of the vapor (m/s)

d

d,,,

Entrained liquid droplets must be eliminated, so that the vapor leaving the still is dry. Therefore, the rising vapor should not exceed a maximum velocity w. Otherwise, liquid droplets will be carried out of the still and mechanical mist eliminators must be used, for example, 0 0

0

simple or finned baffles radial mist eliminators (spinning body with guide vanes) demister (separation mat made of wire or plastic gauze of thickness 100-500 mm and, for the special case of separating small droplets out of a fog, 300-600 mm)

A higher vapor rise velocity w and a decreased free vapor space can then be employed compared with those units without mist eliminators. The vapor rise velocity w is el - e g

liquid mixture empirically evaluated constant, which depends on the type of mist eliminator (C = 0.036 m/s for stills without separation aids, [2.6], C = 0.107 m/s for demisters in regular process operation in columns and stills [2.7])

(2-8)

4 . Dmax. Mg

z,= 71

*

d2

- e,

*

VZ

(2-9)

with vz = 500 +

40

zA5 + 0.01

(2-10)

where vz is the maximum steam space load (m3/(m3 h) and Mg the molar mass of the vapor (vz as given in Eq. (2-10) is valid for stills without a mist eliminator [2.8]). In the distillation period the heat required to heat the contents of the still B, to the boiling point, and the actual heat Q required for the distillation, are provided by the heating system of the still:

where

Qvw, QD heat required to heat the contents of the still and for the distillation process, respectively hBa,hs enthalpy of the still contents B, after charging and at boiling conditions, respectively Ah,, enthalpy of vaporization of the distillate

2.2 Discontinuously and Continuously Operated Simple Distillation

In practice there are different types of heating device available for use on the still [2.9], for example, steam jacket or heating coils of round or half-round tubes welded on the still (since the surface to volume ratio of the still is inversely proportional to the still diameter, this heating system is only used for small heat fluxes or small still diameters) heating coils in the still (mainly for long tubes and large still sizes, although contamination and incrustation possible) heating tube bundle in the still, finned if necessary, (compared with heating coils: less tube length and smaller stills, but still a danger of contamination and incrustation) external reboiler with pump recirculation direct firing Figure 2-6 shows the technical construction of a distillation still, including the steam jacket and the necessary connections.

IV

'1'

VI

Fig. 2-6. Jacketed still. I Adapter, agitator I1 Man hole, sight glass

111 Outlet, distillate vapor

IV V VI VII

Inlet, liquid mixture

Inlet, heating agent Outlet, heating agent Outlet, residue

II

107

2.2.2 Continuously Operated Simple Distillation With continuously operated simple open distillation processes, a continuously fed liquid mixture is partially vaporized by supplying heat to the system. The vapor and residue generated are continuously withdrawn from the distillation unit. Similar distillation devices include vaporizers, which are used for concentrating solutions (see Chapter 7.2.1): circulating evaporator - which has a large liquid holdup and a correspondingly long residence time in the vaporization area; continuous flow evaporator - which has little liquid contents and a short residence time, and therefore offer gentle thermal treatment of the liquid mixture. Continuously operated simple distillation is similar to discontinuous distillation in that it is only of technical interest with large relative volatility values. It is used for mixtures which easily are separated due to large differences in the boiling points of the components of the mixture, but only then the yield of the separation is satisfactory. Continuous distillation is particularly suitable for the separation of small concentrations of low-boiling components from highboiling components, or for the separation of small concentrations of high-boiling components in low-boiling components. Distillation devices are usually continuous flow evaporators. In rectification processes, continuously operated vaporizers are used as simple continuous distillation devices. The vaporizer, usually a circulating evaporator, forms the first practical separation stage of the stripping zone of the rectification column. A mathematical description of continuous simple distillation gives equations similar to those describing discontinuous simple distillation. The concentrations of the liquid mixture and the distillate vapor depend

108

2 Distillation and Partial Condensation

in discontinuous distillation on distillation time, in continuous distillation they are functions of the heat transfer area (see Figs. 2-4 and 2-7). For example, for continuous distillation of a binary mixture, as shown in Fig. 2-8,

the continuous flow evaporator is considered to be a film, or falling film evaporator. The mean distillate concentration along the length of the evaporator tube y D = x, is derived from the mass balance of the vaporizer (2-12) where . . Fa, F, Xa,

I

Z

Evaporator tube length z-

Fig. 2-7. Liquid concentration x, distillate concentration y , liquid temperature tYF and distillate temperature tYD as a function of the evaporator tube length z. Key component of x and y is the low-boiling component (see also Fig. 2-9).

t

xlu

Fa - Fu

=

d

feed and outlet flow rate of the liquid mixture mole fraction of the lowerboiling key component in the liquid phase of feed and outlet flow rate distillate flow rate generated along the length Z of the evaporator tube

The differential equation describing simple continuous distillation is derived from a mass balance for a length element dz of the evaporator tube (Fig. 2-8), giving

d F- d x F y-x TW

(2-13)

Integrating gives

(2-14) l

z

Fig. 2-8. Film evaporation or falling film evaporation in an evaporation tube. HM Heating medium side TW Evaporator tube wall LF Liquid film 0 d Inner diameter of evaporator tube DD Distillate vapors

This is similar to the method already discussed for discontinuous simple distillation. All the assumptions made are also valid. (Under the same conditions, the relationships given by Eqs. (2-4) and (2-5) may also be applied.) The heat flux, dQ, supplied by the surface area element dA = d n dz is

-

e

2.2 Discontinuously and Continuously Operated Simple Distillation

where k is the heat transfer coefficient, and AV the driving temperature gradient between the heating medium and the boiling liquid mixture in the interval of the evaporator tube z and z + dz. (The temperature difference AV is approximately linear if saturated steam is used as the heating medium. If the heat transfer medium is, for example, liquid diphyl, then Ad is the logarithmic mean value.) The heat flux d Q required to vaporize the distillate vapor (with flow rate &) from the boiling liquid film is given by (2-16) where is the vaporization enthalpy of the liquid mixture with respect to concentration and operating pressure at dA. Combining Eqs. (2-13)-(2-16), the required length of the heat evaporator tube 2 is

tors and head or bottom product vapor compression are also discussed. Thin film distillation and rectification may be combined in such a way that the thin film evaporator is connected to a column with both an enrichment and a stripping zone, or just an enrichment zone [2.17]. Such a combination is particularly useful for separation of sensitive mixtures in a vacuum.

2.2.3 Heat Requirement of Simple Distillation Units The heat requirement of a separation unit follows from the energy balance as shown in Chapter 1.3.1. For a simple distillation unit the heat flux may be derived, as shown in Fig. 2-9. An energy balance of the distillation unit under isobaric operation gives Q = D . fiD + F w . fiw - F a

ha = = Ij (fib - hw)+ Fa (Ew - ha) (2-18) *

(2-17) Z is the length of the evaporator tubes required to generate the distillate flux B. If the feed mixture is at a lower temperature than the boiling temperature, a preheating zone is additionally required. For the design of distillation facilities, it is generally sufficient to use A&,,, k and At9 averaged over the length 2 in Eq. (2-17). Methods of calculating the heat transfer coefficient k are found in [2.10]. Applications of thin film and falling film evaporators in film distillation are reported in [2.14], together with the achieved separa-

109

*

*

The specific heat requirement, which is the required heat flux for each kmol distillate vaporized, is Feed Fa,x a

2.

1r

h,

A Distillate vapor D,yo=xh.3o.ho

* * -+

d_

<

Heat supp~y

* -+ -c

F,,

X,,3,.hw

v Outlet,

distillation residue

110

2 Distillation and Partial Condensation Component 1 Mixing unit

(2-19) The specific heat requirement 4 can then be calculated after integration of Eq. (2-14). With the aid of the enthalpy-concentration diagram q arises graphically. MERKEL and PONCHON introduced the caloric enthalpy-concentration diagram. The enthalpy h of a binary mixture is plotted against the mixture composition w with temperature as a parameter. The h, w diagram is shown in Fig. 2-10. If, at a temperature Lp, w kg of component 1 is isothermally mixed with (1 - w)kg of component 2, in a mixing vessel, the enthalpy of mixing Ah, is released. Depending on the temperature and the components in the mixture, the mixing enthalpy may be positive or negative, (i. e., exothermic or endothermic). It is also a function of the composition of the mixture and can change its sign in the concentration range 0 Iw I1. In the MERKELand PONCHON h, w diagram, if the values of the enthalpies h , and 11, of each individual component and the mixing enthalpy AhMhave been found experimentally, the enthalpy of the mixture h can be found from

Component 2 1- w I kg),9, h

*

Mixing enthalpy

Fig. 2-10. Enthalpy-concentration diagram of MERKELand PONCHON. h Enthalpy of the binary mixture w Mass fraction of the key component in the binary mixture

D. Point C is the point on the dew line of composition yo. If the mixture is fed at its boiling temperature, points A and E have the same position and clearly the heat reh ( w ) = w * h, + (1 - W ) . h2 f Ah, (2-20) quired to reach the bubble point temperature is zero. The specific heat requirement 4 for disRepeating the procedure at different temperatures gives a set of isotherms h(w,Lp). continuous simple distillation is derived by The distillation process can be discussed a method similar to that used to determine with the aid of a h, w or h, x diagram. The the heat requirement for continuous simple distillate vapor enthalpies are based on the distillation. For the exact determination of dew line, and the liquid mixture enthalpies 4, the amount of distillation vapor as a are based on the bubble point line of the bi- function of the concentration of the still nary mixture, as shown in Fig. 2-11. The contents and distillates must be considered. specific heat requirement q is therefore the Discontinuous distillation may therefore be distance CDin the 6, x diagram shown in looked upon as a number of differential disFig. 2-12. This line constructed by using tillation processes connected in series. q Eq. (2-19), and by extrapolating the line FA may be graphically determined using the fi, back to the intercept with y,, giving point x diagram [2.11].

2.2 Discontinuously and Continuously Operated Simple Distillation const.

p

p = const

111

I

,Bubble point line

Xw

x.y

3

Liquid phase region

bl

x1

x2 Y1

$01

Xa

YO

__c

Fig. 2-12. Graphical determination of the specific heat requirement q for a continuous simple distillation. A State point of feed B State point of distillation residue C State point of saturated distillate vapor specific heat requirement qv for Line heating the mixture to the bubble point Line specific heat requirement q

Y2

Mole fraction of the low boiling component in the liquid and vapor phase x,y-

Fig. 2-11. fi, x-Diagram of a binary mixture including liquid phase and vapor phase (a). Construction of bubble point line and dew point line in the 6, x-diagram using a boiling diagram (b). A , , A2 and B,, B, state points of liquid phase and vapor phase in equilibrium. 6 Enthalpy x, y Mole fraction of the low-boiling component in the liquid phase and vapor phase Ly Temperature

2.2.4 Flash Distillation In flash distillation, a liquid mixture Fa at its boiling temperature (under pressure) is released into a vapor-liquid separator or a rectification column. No heat is added or removed (isenthalpic throttling, see Fig. 2-13). The liquid mixture may, for example, come from a reaction unit operated at high pressure. It also may be preheated to its boiling temperature in a preheater, operated at a higher pressure than the following separator. Alternatively, the liquid mixture may be heated in a tube heater in which some vapor is generated. The vapor and remaining liquid are in constant contact (“closed simple distillation”).

112

2 Distillation and Partial Condensation

residue

Fig. 2-13. Flash distillation unit. P Preheater S Vapor liquid separator “Expansion or flash evaporator” C Condenser

When a > 1 the released vapor generated under pressure contains a higher concentration of low-boiling components than the mixture introduced into the separator. In flash operation processes, therefore, a partial separation of the feed occurs in which the separator can at best only act as one theoretical separation stage. A mass balance for the separator gives the ratio FJD with the nomenclature as explained in Fig. 2-13. With yo = x,, Fu - x D - x a

D

xa-xx,

x w xa X+

Fig. 2-14. Working diagram for flash distillation. EC Equilibrium curve OL Operating line of slope K (tanx = --FW/B) y Molar fraction of the low-boiling component in the vapor x Molar fraction of the low-boiling component in the liquid

(2-21)

this ratio gives the operating line for a flash distillation, as seen in Chapter 1.8. For a binary mixture, this is shown in Fig. 2-14. If for a given mixture flow rate Fu with mole fraction xu, a vapor flow rate Ij is generated, the operating line can be determined using the operating diagram

(Fig. 2-14). B is the intersection of the equilibrium curve with the line of slope F J D starting at point A, xu. The distillate concentration yo = x, and distillate residue concentration x,, are simply the coordinates of B. Provided that the released vapor and distillation residue are in phase equilibrium, this is approximately true in practice.

2.3 Carrier Distillation

A heat balance over the separator gives the amount of distillate generated during the flash operation.

113

2.3 Carrier Distillation

In carrier distillation, the boiling behavior of liquid mixtures is influenced positively (with respect to a separation effect), by the addition of a n auxiliary component as a vaThis is only worth considering for large por. Figure 1-21 shows the boiling diagram amounts of feed Fa or when the pressure for two immiscible components over the differences on release are large. whole concentration range. Each compoIf some of the distillation residue is recy- nent behaves as though it exists alone; the cled back into the preheater, with flow K, partial pressure in the vapor phase is the then the amount of distillate dkis saturation pressure of each component at the boiling temperature of the mixture and thus the operating pressure is the sum of the (2-23) saturation pressures of each of the components. Due to the immiscibility, the boiling point of the mixture is lower than the boil(In Fig. 2-13, K is the nomenclature used ing points of the pure components at the for the recycle.) operating pressure. For example, the boiling From an energy balance applied to the point of pure o-xylene is 144°C at 1 bar, preheater the heat Q, required in the whereas a mixture of water and o-xylene preheater is boils at 95 "C. The vapor generated during the distillation has a constant composition, which is independent of the composition of the liquid and is an azeotropic mixture. Carrier distillation is used to gently separate nonvolatile impurities from high-boiling substances. Saturated or overheated vapor is blown through the impure liquid in a still. The carrier and the product liquid are The diameter d of the separator can then be immiscible in liquid phase, for example a calculated using Eq. (2-7). The required higher fatty acid-water system. The carrier minimum height ZD of the free vapor space vapor is used firstly as a heat transfer mecan be found using Eq. (2-9), in which vz is dium, heating the contents of the still to the the maximum vapor space loading chosen boiling point and providing the heat for the distillation. Secondly, it is used as a carrier such that for the substance being distilled. The vapor mixture composition corresponds to the vz = 3600 m3/(m3 . h ) (2-25) aceotropic composition. A discontinuously operated carrier distilIn practice, flash distillation is rarely used lation unit is shown in Fig. 2-15. Neglecting as a single separation stage; more often it is the heat due to the nonvolatile impurities, used in combination with other distillation the heat balance over the distillation still, processes without a preheater to arise the gives the required amount of carrier, mostly flash effect. steam DH blown into the still.

(2-22)

114

2 Distillation and Partial Condensation

Fig. 2-15. Steam distillation. DB Still C Condenser cooler PS Phase separator LIC Level indication, control SH Supplementary heating

A = D,Po, B

Po, w

(2-26) where

B

DH fiB ~

AD

b S

quantity of substance to be purified (kmol) total quantity of auxiliary steam (kmol) molar enthalpy of still contents before heating enthalpy of the carrier steam (related to the steam pressure PD) enthalpy of the auxiliary substance at the boiling point (related to the boiling point 19, of the mixture) enthalpy of the vapor of the substance to be purified at Po, B enthalpy of the vapor of the auxiliary substance at po,

W distillation residue saturated vapor pressure of the substance to be purified (related to the boiling point rS, of the mixture) saturated vapor pressure of the auxiliary substance (related to the boiling point t9,)

The boiling point rSs of the mixture can easily be determined graphically using the vapor pressure curves, as shown in Fig. 2-16. The sum of the saturated vapor pressures of both components po,Band po, gives the operating pressure p. P = Po,B -k Po, W

(2-27)

For the azeotropic composition of the distillate, it follows that YE =

Po, B P

(2-28)

~

and the ratio YB --Po,B

--~

w

Yw

Po, w

(2-29)

2.3 Carrier Distillation I-.

2

Q

%P

Operating pressure p

Q

i Pow

d

-1

1

3s

D

115

Fig. 2-16. Determination of the boiling point LS, and the composition of the azeotropic mixture. po,B(19)Vapor-pressure curve of the substance to be purified p-po,w (8) Vapor-pressure curve of the auxiliary substance, substracted from the total pressure p p Total pressure po Saturation pressure c9 Temperature

C

Fig. 2-17. Steam distillation unit with heat recovery by mechanical vapor compression [2.12]. Representation according to the data of GEA Wiegand GmbH. 1 Still 2 Steam converter (inclined tube heat exchanger with vapor condensation in the tubes) 3 Condensate separator 4 Steam compressor (reciprocating compressor with speed regulated direct current motor) 5 Steam saturator A Live steam B Contaminated vapor C Contaminated condensate D Condensatelfeed water E Low-pressure steam F Recovered steam

Data of the presented unit: Steam mass flow feed Steam feed overpressure Performance coefficient (evaporation energy of generated steam feed/electrical power consumption of the compressor)

6.1 t/h

3 bar I

I16

2 Distillation and Partial Condensation

The saturated vapor pressures po,B and p o , w of the substance to be purified and that of the auxiliary substance can be obtained from Fig. 2-16. The dimensions of the still can be determined according to the description in Chapter 2.2. Carrier distillation using water vapor as the carrier is common in the oil and fat industry. In the oil and petrochemical industries, carrier distillation is used to strip loaded solvent from absorption units, using water vapor or an inert carrier in continuous counterflow columns. A carrier distillation unit designed to save energy is shown in Fig. 2-17 [2.12]. Solvent is discontinuously removed from the contents of a distillation still (1) by blowing water vapor into the still. The impure vapor is condensed in a steam converter (a tube heat exchanger (2), slightly inclined). The condensate is fed to a regeneration unit for further treatment. Pure steam generated at the jacket during vapor condensation is compressed to the injection pressure in a steam compressor (4) and is simultaneously superheated. The superheat is removed in a saturator (5) by the injection of condensate or water and the saturated steam is injected into the distillation unit (1). A general discussion of the economic efficiency of mechanical vapor compression is found in [2.13]. The use of vapor compressors is already of economic benefit if

15Yo of the primary energy is saved. For example, the heat recovery, by vapor compression of live steam at a pressure of 5 bar gives a primary energy saving of up to 22 Yo, as illustrated by a diagram in [2.13].

2.4 Vacuum and Molecular Distillation Vacuum distillation is suitable for the separation of thermally sensitive liquids, being a gentle distillation process which operates at low pressure. Due to the reduced operating pressure, the boiling point of the mixture is also reduced. A low residence time and a narrow residence time distribution are maintained, by careful selection of the distillation unit to be used. The advantages of vacuum distillation over ordinary distillation at normal or higher pressure are listed in Table 2-1. Different vacuum ranges, including the vacuum forming device are discussed in Table 2-2. Common distillation units are also shown. Molecular Distillation Molecular distillation is used to separate very temperature sensitive liquid mixtures at absolute operating pressures in the range 0.1 -0.001 Pa.

Table 2-1. Advantages of vacuum distillation over normal pressure distillation.

0 0

A lower and only short-term heating is nessecary due to the lower boiling point of the mixture to be separated Increase of selectivity due to the increasing relative volatility with decreasing operating pressure A favorable influence on the azeotropic point of a n azeotropic mixture under lowered operation pressure Usage of cheaper heating medium due to a lower temperature level of the process

117

2.4 Vacuum and Molecular Distillation

Table 2-2. Vacuum range, vacuum forming device, distillation process and features of distillation units.

Vacuum range

Vacuum forming device

Distillation process

Distillation unit

Low vacuum (1 bar1 mbar)

Positive displacement Pump Booster diffusion vacuum pump (particularly steam jet pumps)

Low vacuum distillation

Vacuum circulation evaporator Continuous evaporator with falling or wiped film Rotary evaporator

Medium vacuum (1 mbar0.1 Pa)

Pump combination for example: positive displacement pump+ booster diffusion pump (oil ejector)

Medium vacuum distillation

Continuous evaporator with agitated film (film evaporator with fixed or swinging wipers, film evaporator with brushes or rolls) Rotary evaporator

High vacuum (0.1 Paca. 0.001 Pa)

Pump combination for example: two stage positive displacement pump + diffusion pump or single stage positive displacement pump +oil ejector +diffusion pump

Molecular distillation (open path, short path distillation)

Cup molecular distillator Rotary molecular distillator Film molecular distillator Centrifugal molecular distillator

The maximum evaporation rate is obtained if all molecules distil from the liquid into the vapor space and reach the condensation unit without collision with other molecules. After condensation the liquid is withdrawn from the unit. Distillation without recondensation effects, such as reflection of vapor molecules back to the liquid phase, or time delays due to collision with other molecules on the way to the condensation unit, is the aim of molecular distillation. The mean free path of the molecules at a pressure of 0.1 to 0.001 Pa is a few centimeters, then the distance between the vapor-

ization surface and the condensation surface should be less or approximately this mean free path of the molecules in the distillation unit. Vaporization from the surface occurs without vigorous agitation of bubbles, unlike the case of nucleate boiling, where bubbles are formed and agitated at the heating surface. The liquid mixture spreads over the heating surface becoming a thin film of constant thickness. The rate of the distilled vapor flow is derived from the kinetic theory of gases. this According to KNUDSENand LANGMUIR, flow rate is

aA

2 Distillation and Partial Condensation

118

dA= E

*

1.577 . lo6 -p0,;

vzT

(2-30)

where p0,;is the saturated vapor pressure (given in bars), i is the substance to be distilled at temperature T with molar mass Mi. dAis obtained in kg/(m2 h). E is the effectiveness factor, E 5 1. Its value increases the fewer vapor molecules diffuse back, and the better the renewal of the active liquid film surface. With ordinary distillation the relative volatility is given by Eq. (1-144). For the pressure range of molecular distillation, not only are the saturation pressures p , ; and po,jimportant, but also the molar masses of the components i a n d j . The relative volatility is thus

.I

-

-.,I

.

.

(2-31)

During equilibrium distillation each component vaporizes proportionally to its partial pressure p i ,pJ, etc. During molecular distillation each component is vaporized proportionally t o p , / pJ / v M J , . . The composition of the distillate X i , for component i, distilled from a binary mixture of components i and j , is given by

mL,

-5

B

4, 4 7

Fig. 2-18. Agitated film evaporator for molecular distillation. 1 Agitator with rolls or brushes to circulate the

film 2 Feed 3 To vacuum pumps 4 Heating jacket; heating agent in, out 5 Condenser (inlined) 6 Cooling agent in, out 7 Distillate

8 Residue

(2-32) Investigation and operating costs make molecular distillation an expensive separation method. It is used mainly in the pharmaceutical industry to treat high-grade substances such as vitamins, hormone concentrations, fats, oil, and waxes. Agitated thin-film evaporators with a falling, revolving liquid film (Fig. 2-18) are often used in molecular distillation. The film is 0.1-1 mm thick, with a mean residence time ranging from 10 s-10 min and a

throughput rate of up to 1 t/h. A further device is the centrifugal molecular distillation still (Fig. 2-19). It has a film thickness of 0.03-0.001 mm depending on the number of revolutions of the rotor. A typical throughput is 1 t/h, with a mean residence time of 0.1-0.001 s [2.15, 2.16, 2.18-2.20, 2.1401.

2.5 Countercurrent Distillation (Rectification) 5

6

Fig. - 2-19. Centrifugal - molecular distillation still. Rotating evaporator with liquid film forming and mixing due to the effect of centrifugal force Electrical radiant heater Condenser with cup surface and drainage to collect the distillate (aerofoil condenser with self-pumping effect for distillate) Cooling agent for condenser Distillate discharge To vacuum pumps Feed Residue discharge

2.5 Countercurrent Distillation (Rectification) Countercurrent distillation or rectification was introduced as a special countercurrent distillation process in Chapter 2.1. With the requirement of high quality thermal separation, the reflux flows down while vapor flows up the rectification column, thereby ensuring intensive contact of the phases. Intermediate products from, for example, a reaction unit, usually need treatment in an energy intensive thermal separation process to obtain the desired purity of the final product. 60 to 80% of the energy used in the chemical industry is spent for this purpose [2.21]. An energy intensive rectification process for thermal separation is of-

119

ten employed, and thus the required rectification columns are designed and linked to recover and save as much energy as possible. An analysis and optimization procedure for the energy usage in each individual unit and the entire plant, as introduced by Linhoff [2.21 a] is often used. In the following sections, different rectification processes and methods for the design of rectification columns are discussed. Energy saving possibilities are highlighted.

2.5.1 Process Variations of Rectification A number of variations of the rectification process have been developed and adapted to suit certain rectification problems during the decades of distillation usage. Substantial adaptations are due to the properties of the mixture and its thermal behavior, the mode of operation, the required operating conditions, the economic efficiency, the flexibility, and the safety of the operation. Several important process variations are now presented and briefly discussed. 2.5.1.1 Continuously Operated Rectification in Rectification Columns with Enriching and Stripping Zones

Figure 2-2 shows the most common process in rectification practice. It can give a good separation efficiency for nonazeotropic mixtures, depending on the number of separation stages and the reflux ratio. It is possible to almost completely separate a nonazeotropic binary mixture into its pure components in a column including enriching and stripping zones. To separate a multicomponent mixture, several columns must be linked. For a nonazeotropic mixture with k components in the mixture, the

120

2 Distillation and Partial Condensation

For example, to separate a mixture of four components, 10 different separation tasks for 10 different mixture streams with 5 different column arrangements in which each arrangement contains 3 rectification columns are possible (Fig. 2-20). During optimization of the separation process (minimizing the energy and fixed costs), additional variations of Z,, may be included, as shown for a ternary mixture in Fig. 2-21 [2.24, 2.251.

required number of columns is k - 1. The possible number of different connections of the columns, Z, (variations of separation) is given in [2.22, 2.231. Provided each column performs only one separation task with no side streams, recycle, and circulation - l)] ! z,= [2k ! (k (k-l)! *

(2-33)

-

*

The number of mixture streams is

z,=-k2

(2-34)

* ( k +1)

and the number of separations Z, (separation tasks) of the mixture streams is

z,= k6 . (2- 1)

(2-35)

-

2.5.1.2 Stripping (Exhausting) Column Figure 2-22 shows a stripping column operated without an enrichment zone and without reflux. In this variation, a good purity of bottom product can be achieved with a small throughput. This variation is of interest for small amounts of overhead product,

iu C)

L

3

2

4

L

3

Fig. 2-20. Different methods of connecting three columns in a quaternary mixture distillation unit. a)-d) Serial connections of columns e) Parallel connection F Feed 1, 2, 3, 4 Components in the order of increasing boiling points

2.5 Countercurrent Distillation (Rectification)

$i2

121

3

’ 3

or for the recycling of overhead product, or for further treatment of column sidestreams (Fig. 2-25 shows sidestream strippers for a crude oil column). Since the number of separation stages is fixed the only operation parameters are the bottom temperature and the vapor loading, which is ad-

-

Fig. 2-21. Methods of connecting columns in a ternary mixture distillation unit, without vapor recompression. (Representation according to SCHLUTER and SCHMIDT [2.24], and BDDER [2.25]).

justed by the heat transfer per unit surface area in the vaporizer.

2.5.1.3 Enrichment Column Figure 2-23 shows an enrichment column in which no reboiler is used. The feed to C

A

Fig. 2-22. Charge reflux fractionator (stripping section only). SC Stripping column R Reboiler C Condenser F Feed, liquid E Overhead product A Bottom product

Fig. 2-23. Charge reboiler fractionator (enrichment section only). RC Rectifying column C Condenser F Feed, saturated vapor or wet vapor E Overhead product R Reflux A Bottom product

122

2 Distillation and Partial Condensation

C)

K P H

PS &A I

E --GLS

T

o-& T

% /jc

the bottom must therefore be saturated vapor, or at least wet vapor, with a small fraction of the liquid phase. With a small throughput, the overhead product is of good purity. With a fixed number of separation stages, the reflux ratio is the only operating parameter. A column operating only in the enrichment mode is used with carrier distillation in crude-oil distillation

Fig. 2-24. Carrier distillation, “steam distillation”. a)-c) Different types of steam distillation d), e) Regeneration of loaded solvent from an absorption stage, stripping RC Rectifying column SC Stripper C Condenser R Reboiler PS Phase separator GLS Gas-liquid separator PH Preheater CC Condenser cooler F Feed B i Rich oil, solvent + absorbent E Overhead product i‘ Carrier, water T b Vaporous carrier, steam or inert gas, stripping in solvent regeneration A Bottom product RL Recovered solvent, lean oil

processes. This variation may also be used if the feed mixture shows a tendency to form a scale (incrustation), or if small streams of heavy boiling components are diverted in repeated rectification stages. The enrichment column now used with a circulating evaporator and a receiver is operated in the “endless circulation” experimental rectification process.

2.5 Countercurrent Distillation (Rectification)

2.5.1.4 Carrier Rectification Carrier rectification can be used with different variations, to gently separate highboiling substances from a mixture, as shown in Fig. 2-24. Carrier steam which is immiscible with the components in the mixture is introduced into the column. Since the partial pressure of the components of the mixture is reduced by the additional vapor component (steam), the boiling point of the mixture is reduced. The carrier steam can also be used as heating medium (see Chapter 2.3). 2.5.1.5 Combinations of Different Variations The discussed variations may be combined, depending on the number of mixture components and their boiling behavior, to form rectification units consisting of many individual columns, the aim being optimum economic efficiency. Figure 2-25 shows a

123

crude-oil column with sidestream strippers; the first column, an enrichment column, is used to separate the crude-oil, with carrier rectification used in the sidestream strippers. The sidestream strippers act only as stripping columns, recycling the overhead product back to the main column. 2.5.1.6 Rectification with an Entrainer

More theoretical stages, and a reflux ratio which increases with time, are needed to separate a liquid mixture by rectification the closer the relative volatility ai,, approaches to unity. A simple separation is not possible if ai,j= 1. According to COLBURN and SCHOENBORN the number of theoretical stages required is 4 / ( a , ,- 1) for a 99% separation. For ai,j = 1.05 80 theoretical stages are required, while for ai,j = 1.1 under the same operating conditions only 40 are required. An entrainer or auxiliary component must be added to mixtures with small dif-

w

Fig. 2-25. Crude-oil distilla-

CG

C

tion unit. AC Atmospheric column vc Vacuum column HE Crude-oil heater, directly fired C Condenser PS Phase separator ss Sidestream stripper 0 Crude oil Steam w Water CG Crude gasoline K Kerosene G Gas oil H Heavy oil v Vacuum gas oil R Residue

s

124

2 Distillation and Partial Condensation

ferences in the components’ boiling points (ca. a < 1.04) and to azeotropic mixtures to influence the liquid phase volatility behavior, and to shift the location of the azeotropic point (see also azeotropy in Chapter 1.4.3.2). Using an entrainer for distillation is only economically efficient for mixtures with nonisomeric components. Azeotropic rectification or extractive rectification is used depending on the behavior of the mixture. An extractive rectification (Distex-process, distillation extraction) can be used to separate substances with similar chemical properties and small differences in boiling points. By reducing the partial pressure or the volatility of one of the components, the adjuvant increases the relative volatility of the original binary mixture over the entire concentration range. Azeotropic rectification can be used where there is a low concentration of the key component in the feed. Entrainer and key component are discharged as overhead product. Since the quantity of the entrainer is small, the additional energy required due to its presence is small. The entrainer is chosen such that it forms an azeotropic mixture with

either one of the mixture components, having a minimum boiling point due to the partial miscibility in the liquid phase. (A comparison of the economic efficiency of rectification processes, including auxiliary substances, can be found in [2.26].) Choosing the entrainer to separate a certain mixture is a substantial problem. Properties of mainly polar entrainers (used in azeotropic and extractive rectification) are given in Table 2-3. Of the possible interacting forces between the liquid molecules of the mixture, hydrogen bonds are most important. A classification of important groups of substances with respect to the occurrence and strength of the hydrogen bonds has been done by BERG[2.27]. These groups are listed in Table 2-4. A general overview of expected deviations from Raoult’s law for mixture substances is given in Table 2-5, in the classes I-V. Thus first pointers toward choosing a rectification entrainer are given. Experiments with different entrainers are necessary to make a final choice. The selectivity S C j is the decisive factor

Table 2-3. Required properties of an entrainer to be used in azeotropic or extractive rectification. 0 0 0

0 0 0 0

0 0 0

Suitability to cause a change in the relative volatility ai,j of the original mixture components i, j High efficiency and selectivity (a small amount of entrainer should cause a suitable change in a l j ) Total miscibility with the mixture components at the operating temperature with extractive rectification, and with azeotropic rectification formation of a miscibility gap and a minimum-boiling azeotropic Easily separable from the mixture components Low evaporation enthalpy Thermal stability No toxicity, low environmental pollution Favorable physical-chemical properties (e. g., boiling point, melting point, viscosity, surface tension, density) Noncorrosive Good availability, low price

2.5 Countercurrent Distillation (Rectification)

125

Table 2-4. Classification of important groups of substances. Liquids which form a three-dimensional network of strong hydrogen bonds, *, ** Examples : water, polyhydric alcohol, aminoalcohol, hydroxylamines, hydroxyacids (oxyacid), amides, polyphenol I1 Liquids whose molecules contain active hydrogen atoms as well as donor atoms (oxygen, nitrogen, fluorine) *, ** Examples: alcohols, acids, phenols, amines, oximes, ammonium, nitro compounds with a-state hydrogen 111 Liquids whose molecules contain donor atoms but no active hydrogen atoms** Examples: ethers, ketones, aldehydes, esters, tertiary amines, nitro compounds without a-state hydrogen, nitriles IV Liquids whose molecules contain active hydrogen, but no donor atoms Examples : chloroform, dichloroethane V All other liquids Examples: hydrocarbons, carbon disulfide, sulfides, mercaptan

Class 1 Class

Class Class Class I

* **

Associated liquids covered by class I and 11. Water-soluble liquids covered by class I and 111.

where is the relative volatility, K?and K? are the equilibrium constants of the key components i and j , with activity coefficients yi and y j , respectively. p and a indicate whether the adjuvant is present or not. The larger the value of the selectivity Si,j the better the efficiency of the adjuvant in the rectification process. Table 2-6 gives an explanation of how the adjuvant normally functions and how it affects the relative volatility as a result of this. Azeotropic Rectification

In azeotropic rectification, the otherwise unwanted formation of an azeotrope is used to simplify the separation by distillation of a mixture with a narrow range of boiling points, or azeotropic mixtures. An auxiliary

component chosen to form an azeotropic point (in mixtures with a narrow range of boiling points) or to change the azeotropic point (of azeotropic boiling mixtures) is added to the mixture. Since the newly formed azeotrope is low-boiling, it can be easily withdrawn from the column as an overhead product. In contrast to extractive rectification, the boiling point of the entrainer should be similar to those of the components of the mixture (temperature range 35°C). Entrainers forming a heteroazeotrope are preferred, as they are easily separated mechanically. Azeotropic rectification processes using adjuvants to form a homogeneous azeotrope are not of economic interest, since distillation is not easy. Therefore, other separation processes are employed. Figure 2-26 shows an azeotropic distillation unit [2.5, 2.28-2.301 for a binary mixture. In an azeotropic column, a binary mixture is separated into an almost pure bottom product component 1, and a low-boiling ternary mixture as an azeotropic over-

*

126

2 Distillation and Partial Condensation

Table 2-5. Deviation from Raoult’s law of mixtures classified in Table 2-4. Liquid mixtures of classes

Deviation from Raoult’s law

I+V

I1 + v

Positive (at I + V often miscibility gap)

I11

Negative

+ IV

I + IV I1 + IV

Positive (at I + IV often miscibility gap)

I+I I + I1 I + 111 I1 + I1

Usually positive, sometimes negative formation of an aceotropic with maximum boiling point

111 + I11

Almost ideal behavior, non- or only small positive deviation, seldom formation of an azeotropic with minimum boiling point

I1 + 111

I11 + v IV + IV IV + v

v+v

Table 2-6. Effect of entrainer on the mixture behavior. Effect of entrainer

Consequence

Formation of a complex or a hydrogen bond with the key component i

Lowering of the vapor pressure of component i Decrease of the relative volatility ai,j

Complex dissolving of components i, j or dissolving of associates of similar molecules

Increase of the relative volatility

Formation of an easily separable azeotrope with key component i

Influence on the relative volatility ai,j

head product. After condensation of the overhead product, two immiscible phases form which are decanted in the phase separator. The light phase, with a high concentration of the entrainer 3, is the reflux to the azeotropic column. The heavy phase is separated into an almost pure component 2 and the entrainer in the entrainer recovery column, which is operated as a stripping column.

Other columns arrangements than those shown in Fig. 2-26 are possible in azeotropic rectification processes. The column arrangement and the locations of the outlets for the components of the mixture and the entrainer, depend on the state and behavior of the mixture. Examples of practical azeotropic rectification are listed in Table 2-7.

2.5 Countercurrent Distillation (Rectification)

MC

,

RC

Fig. MC RC R C PS F 3

$3

127

2-26. Azeotropic distillation unit. Azeotrope tower (main column) Entrainer recovery tower Reboiler Condenser Phase separator Mixture feed of components 1 and 2 Entrainer Entrainer make-up

Table 2-7. Examples of a practical azeotropic rectification process. Mixture to be separated

Entrainer

Ethanol - Water (homogenuous azeotrope)

Benzol, Trichlorethylene

Water - Acetic acid (not an azeotropic mixture although costly distillative separation without an entrainer)

Ethylene dichloride, n-Propyl acetate, n-Butyl acetate

Acetone

- Methanol

Aromatic-rich fraction in crude oil rectification (separation of the remaining paraffines)

Extractive Rectification

Extractive rectification (Distex-process) [2.28, 2.29, 2.35, 2.361 is a distillation process which separates homogeneous, narrow boiling, or azeotropic liquid mixtures by the aid of an added substance of low volatility. The boiling point of the entrainer must be substantially higher than the boiling points of the components of the mixture. The less volatile component of the mixture selectively bonds to the entrainer, due to interactive forces. The vapor pressure of the less volatile component is therefore reduced and

Methylene chloride Methylene ethyl ketone

hence the relative volatility of all the other components are increased. Separation of the lower-boiling components thereby becomes easier. The activity coefficient, and hence the partial pressure and volatility, of only one component should be changed by the adjuvant. For example, adding aniline to a mixture of benzol and cyclohexane reduces the volatility of benzol. This gives the equilibrium curve a more favorable position with respect to separation efficiency. A high selectivity, or effectiveness of the entrainer leads to large heats of mixing (of

128

2 Distillation and Partial Condensation

either sign) in the mixture. Gas chromatography may also be used to evaluate the selectivity, and the behavior of the components in the mixture in the presence of the added component [2.31]. A methodical procedure to choose the component to be added is given in [2.37]. This is based on the Unifac method, which derives the phase equilibrium behavior of the mixture from the molecular structure of the mixture components. Figure 2-27 shows a setup of an extractive distillation unit. The adjuvant 3 is added at the top of the extraction column and withdrawn with the less volatile component 2. Both compromise the bottom product. This is then separated into the more volatile component 2 and the adjuvant 3 in a solvent stripper. This is then recycled back to the extraction tower. Table 2-8 lists some examples of practical extractive distillation processes. For the separation of a narrow boiling or azeotropic binary mixture as in azeotropic

Fig. MC RC C R F 3 83

2-27. Extractive distillation unit. Extraction tower (main column) Solvent stripper to recover entrainer 3 Condenser Reboiler Feed of components 1 and 2 Solvent entrainer Solvent make-up (entrainer make-up)

Table 2-8. Examples of practical extractive rectification processes. Mixture to be separated

Entrainer

Benzol - Cyclohexane

Aniline

Methylcyclohexane

- Toluene

Aniline

Hydrochloric acid - Water

Concentrated sulfuric acid

Nitric acid - Water

Concentrated sulfuric acid

C,-Fraction

Furfurol, dimethyl formamide N-methyl pyrolidine N-methyl pyrolidine (NMP) (Distapexprocess, Lurgi [2.33]) N-formyl morpholine (NFM) (Morphylane-process), Morphylex process including a combination of extractive rectification and liquidliquid extraction, Krupp-Koppers t2.341

Pyrolisis benzene, reforming products and other aliphatic/ aromatic mixtures (aromatic substance production [ 2 . 3 2 ] )

2.5 Countercurrent Distillation (Rectification)

rectification two columns are required for extractive distillations. The same arguments also apply to investment and operating costs. The application of these processes therefore always depends on the result of the efficiency calculation. A liquid-liquid extraction process with downstream solvent recovery is frequently an alternative process. The azeotropic mixture of tetrahydrofuradwater is processed by extractive rectification with glycol as the third component; whilst with furan as a third component, it is processed by azeotropic rectification and with glycerine as a third component it is processed by liquid-liquid extraction. Figure 2-28 shows the DISTAPEX-process [2.33] as an applied example of extractive rectification used to separate a mixture of aromatic and nonaromatic substances.

-

129

2.5.1.7 Heteroazeotropic Rectification [2.381 An azeotropic mixture with a widemiscibility gap can be separated by distillation without using a third component. Figure 2-29 shows the separation of a binary mixture with a miscibility gap. In the first column, the feed is separated into an almost pure less volatile component 2 as the bottom product, and into a mixture of nearly azeotropic composition as the overhead product. After condensation, the overhead product decays into two liquid phases by phase separation. According to the miscibility gap, the composition of the phases is xF, and xF2. Phase I, of composition x, is the reflux in the first column, while phase I1 of composition x,,, is fed to the second column. The second

r 4 Li

Extractive distillation

1

Feed

---

EC

Nonaromatic substances

ST

Aromatic substanc s

aromatic substances

L ~

T

Solvent

Fig. 2-28. Lurgi-DISTAPEXprocess to separate aromatic substances from a nonaromatic/aromatic mixture by extractive distillation. Representation according to Lurgi, FrankfurUMain [2.33].

EC Extractive distillation column to separate a mixture into the nonaromatic substances (top product) and aromatic substances + solvent (bottom product) ST Stripper to separate the solvent from the solvent/aromatic substance mixture Examples: Separation of benzol and toluol from pyrolysis benzene; separation of carbonization benzol and C,-aromatic substances from the reformate. Solvent: N-formyl pyrolidine.

130

2 Distillation and Partial Condensation

column is operated in stripping mode, the bottom product being almost pure, more volatile component 1 of the mixture.

b)

a)

2.5.1.8 Two Pressure Operation

The location of the azeotropic point depends on pressure, its position in the equilibrium diagram shifts to the right with a decrease in pressure. A higher fraction of more volatile component is then found in the azeotropic mixture. The separation of such a liquid mixture by distillation is based on this fact. Components are completely miscible and form a minimum or maximum azeotrope without a third component. In Fig. 2-30, this rectification separation process, in two columns operated at two different pressure levels, is explained as a two pressureprocess for a binary mixture. The binary mixture consists of components 1 and 2, with mole fraction x, of the low-boiling component 1. In the first column, operated at a lower pressurepcl, the binary mixture is separated into component 2 as the bottom product, and an azeotropic mixture of composition xAl,as an overhead product. In the second column, operated at a pressure pG2> pG, the azeotropic mixture is separated into component l (at the bottom) and azeotropic mixture xAz (at the top). The azeotropic mixture of the second column is then fed into the side of column 1 at an appropriate location. The two pressure process is particularly economic if the reboiler of the first column is used to condense the top vapor of the second column, with the overhead product acting as the heating medium. This operating mode is possible if there is a large enough difference in the pressure levels of the two columns and if the boiling points of both components 1 and 2 are not too far apart.

+

=2

=I

Fig. 2-29. Azeotropic distillation unit for heteroazeotropes (a) with corresponding vapor-liquid equilibrium diagram (b). RC1, RC2 Rectification columns C Condenser R Reboiler PS Phase separator F Feed of component 1 and 2 Liquid phase 1 with a mole fraction PH 1 xF1of component 1 PH 2 Liquid phase 2 with a mole fraction x,, of component 2 A Azeotrope xF1,xF2 Mole fraction of low-boiling component in both liquid phases in the separator Mole fraction of the lower boiling X component in the liquid Mole fraction of the lower boiling Y component in the vapor

2.5 Countercurrent Distillation (Rectification)

131

2.5.1.9 Diffusion Distillation [2.39]

RC2 PGZ

RC1 oG1

Fig. 2-30. Two-pressure distillation unit with corresponding vapor - liquid equilibrium diagram. RCI, RC2 Rectification columns, operating pressure PG2 > p C l Reboiler R Condenser C Feed consisting of components 1 and F 2 with a mole fraction x, of component 1 Aceotropes A , , A2 Aceotrope mixtures x A l ? xA2 Mole fraction of the low boiling X component in the liquid Mole fraction of the low boiling Y component in the vapor ~~~~

~~~~~

The azeotropic composition of a binary system is not only influenced by recycled third components added to the mixture and by the operating pressure, but also by the presence of an insoluble inert gas, which is favorable for the separation if the volatile components of the mixture diffuse at different velocities across the regions of inert gases. Figure 2-31a shows an example using the binary isopropanol/water mixture. Steam diffuses faster across an air bolster than isopropanol vapor, and thus, the distillate of isopropanol mole fraction x, is more water-rich due to the presence of air between the liquid and the condensation surfaces (point A2), compared with the distillation under air exclusion (point A,). The shifting of A, +A2 is of practical use in separating water rich isopropanol/water mixtures (Fig. 2-31 b). A feed is almost completely separated into water flow rate F$’ leaving column DCl and a isopropanol flow rate P leaving column DC2. Both columns are operated at ambient pressure. The overhead products k,and K2 are practically of azeotropic composition (point A,, Fig. 2-31a). In the condenser c, most of the condensed overhead products form the reflux 6, or R2, to the DC1 Or DC2, The remainder is then separated into a water rich fraction b, and an alcohol rich fraction D7 in the diffusion distillation unit DA (Fig. 2-31 c shows the diffusion separation unit). Some of the overhead product fluxes kl and K2 are heating the unit while water is used for cooling. For the practical design of a diffusion separation unit a tube bundle apparatus consisting of vertical double pipes may be used. The outer pipes are heated while inner pipes are cooled. In the annulus, two falling films, separated by an air gap, flow downward. Special attention should be paid to the generation of the falling film, which

132

2 Distillation and Partial Condensation

%

Y

50

0 Water

XA

x-

100 % lsopropanoi

t

rich

Cold w a t e r

t

outlet

,Alcohol pool

should be stable, with no waves and an inert gas gap as narrow as possible. Convection currents should be considerably oppressed. 2.5.1.10 Overpressure, Low Temperature, and Vacuum Rectification The properties and behavior of the mixture sometimes mean that operation at over-

Fig. 2-31. Diffusional distillation, binary system isopropyl alcohol/water. Representation according to SCHLUNDER [2.39]. a) McCabe-Thiele diagram EC Equilibrium curve CC Diffusional distillation concentration curve b) Diffusional distillation unit DC 1, DC2 Distillation columns R Reboiler C Condenser DA Diffusion separation apparatus c) Diffusion separation apparatus, schematic

pressure or vacuum is necessary. This leads to an increase in the operating and investment costs, compared with the otherwise more favorable normal pressure operation. Rectification under high pressure is used for vaporous mixtures having overhead products of high volatility. Overpressure is also required for the two pressure process. Extremely low-boiling substances such as air, natural gas, and liquid gases are sepa-

2.5 Countercurrent Distillation (Rectification)

133

A--I 1

Bottom product

Fig. 2-32. Vacuum rectification unit with condenser/evaporator, steam jet aspirator, and recirculation of the condensate. Representation according to GEA Wiegand GmbH [2.58]. RC Rectification column R Evaporator C Condenser/evaporator CH Head product cooler SJ Steam jet aspirator CV Condensate vessel C P Condensate pump

134

2 Distillation and Partial Condensation

rated by low temperature rectification. Instead of overpressure, cooling of the mixture components in a liquefying cycle using the JouleThompson effect is carried out in connection with a rectification process. If there are thermally sensitive substances in the mixture, vacuum rectification is necessary and the additional investment and operation costs of the vacuum plant must be taken into account. A vacuum column has a larger diameter for the same mass flow, with the consequence of greater strength, and more costly design (as a first approximation, the cross-sectional area is directly proportional to the vapor mass flow and inversely proportional to the operating pressure). To guarantee a small pressure drop in the vapor flow along the column length, internals such as packing, braided metal and rotating equipment, with a low specific pressure drop are used. Figure 2-32 shows a simplified setup of a vacuum rectification process including a condenser/evaporator, a steam jet aspirator, and a condensate recycle.

follows. The concepts described in the discussion may also be used to estimate the column dimensions required for rectification of multicomponent mixtures with two chosen key components having similar boiling points. To determine exactly the dimensions of a column required to separate binary and multicomponent mixtures, the calculation methods for counterflow columns, shown in Chapter 1.9.3, can be applied. Figure 2-33 shows a rectification unit used to separate a binary mixture consisting of components 1 and 2. x, is the mole fraction of the lower-boiling component 1 in the feed, 8 x, the mole fraction in the overhead product E, and x, the mole fraction in the bottom product A. The overhead product is almost pure component 1 (xE+ l), the bottom product is almost pure component 2 (xA 0). E,, f i E and hA are the molar enthalpies for the appropriate temperatures and compositions of the flow rates, 8 E and A. +

2.5.2 Continuous Adiabatic Rectification In continuous rectification a liquid mixture is continually fed into the separation column. This feed is separated into a continuously withdrawn overhead product, a bottom product, and perhaps side products. The flow rate and composition of the feed and also the required compositions of the product fractions are usually given in rectification problems. The operating conditions and dimensions of the column must be chosen to obtain the desired quantity and purity of the product fractions. A discussion of the continuous rectification of a binary mixture in a column which includes stripping and enrichment zones

Fig. 2-33. Rectification unit for binary systems.

EC Enriching column SC Stripping column C Condenser R Reboiler

2.5 Countercurrent Distillation (Rectification)

135

2.5.2.1 Flow Rates An overall mass balance of the column (Fig. 2-33) for the lower-boiling component 1, gives the product flow rates and, if no sidestreams are withdrawn. as where, if the dependence of the vaporization enthalpies on the pressure is neglected, C is a constant.

c= With known vapor flow G and reflux R the internal mass flow rates may be calculated for any cross-sectional area of the column operated under adiabatic conditions. LS refers to the lower-boiling component and SS to the higher-boiling component. Figure 2-34 shows a height element dz of a rectification column. Along dz the lowerboiling component (LS) is enriched in the vapor flow G from y to y + dy and G becomes G + dG, with either a positive or a negative change. For cases of relative volatility c ~ ~ 1~ this , ~enrichment ~ > of LS in the vapor is due to the condensation of a differential amount dSS of SS out of the vapor. Under adiabatic operation of the column, the heat released dSS Afil,g,ss is used to vaporize a differential amount dLS of LS from the reflux and

A 4,g, ss AhI,g,SS - G , g , L S

(2-41)

-

A mass balance for the lower-boiling component in the vapor phase over the height element dz, as shown in Fig. 2-34 gives G - y + d L S = ( G + d G ) . ( y + dy)

(2-42)

or d S = d G - y + G-dy

(2-43)

and C . dG = d G * y + G * dy

(2-44)

-

G+dG, Y + dY

R +d R x+dx

T-

(2-39) where Afil,g,Lsand Ahl,g,SSare the molar vaporization enthalpies of the lower- and higher-boiling components, respectively. The heat of mixing and heat of superheating can be neglected. > Afil,g,ssthen For example, if d i S < dSS, which results in a decrease in both vapor flow rate and reflux from the bottom to the top of the column. The change in the vapor flow rate dG is

/

‘i __

G, ’

Y

. R’ X

Fig. 2-34. Differential height element of a rectification process. Vapor phase 1 Reflux (liquid phase)

136

2 Distillation and Partial Condensation

If the product d G . dy is neglected, rearranging gives d 6 - dy G c-y

(2-45)

The relationships used to calculate the vapor fluxes and refluxes for the enrichment and stripping columns, based on Eq. (2-46), are listed in Table 2-9. In normal operation, heat losses from the columns are unavoidable. Thus vapor condensation causes an increase in the amount of reflux.

Considering the balance area for a part of the column, between the two cross-sectional areas I and 11, shown in Fig. 2-35, with the corresponding vapor flow rates GI and GII 2.5.2.2 Heat Requirement of a Column of mole fractions yI and yrI, integration of The heat flow Q required to operate the colEq. (2-45) gives umn adiabatically is provided via the reboiler. A heat balance over the rectifica(2-46) tion column shown in Fig. 2-33, gives Additional balance equations from Fig. 2-34, are

GII + 6 , = G I

+ d,,

(2-47)

Q=QC+E*&E+A

(2-49) For the case of total condensation of the overhead product in the condenser and considering the reflux ratio v

R

V = Y

E

6,

I -.

n,

6,

(2-50)

the heat flow Qc from of the condenser is

I

Yn .

‘&A+QV-F*&F

-I

I

.-

II

Fig. 2-35. Balance space for the derivation of Eqs. (2-46-2-48). Vapor phase 1 Reflux (liquid phase)

is the vaporization enthalpy referred to the pressure and composition x, in the condenser. (If the amount of overhead vapor condensed is the same as the reflux to the column, the condenser acts as a dephlegmator. The dephlegmator and the additional condenserkooler required to condense and subcool the overhead product E must be included in the heat balance.) Heat losses Q v from the rectification unit depend on the type and thickness of the insulation material. The reflux is increased by heat losses from the rectification column by a so-called “wild reflux”. These unavoidable heat losses relaxe the concept of “adiabatic” operation.

2.5 Countercurrent Distillation (Rectification)

137

Table 2-9. Relationship between the vapor and reflux flow streams in stripping an enriching column *. 0

Vapor flow G , , at the top of the enriching column above the enriching tray or above the packing retaining grid G,, = E . (v, + 1) Top product mass flow; v, reflux ratio E

0

Vapor flow G,,, at the bottom of the enriching column above the feed tray

y K w= x, Molar fraction of LS in the product (with total top vapor condensation) YKa 0

Molar fraction of LS above the feed tray

Reflux flow R , , in the top of the enriching column

R,,=v;E 0

Reflux ratio vv in any cross section of the enriching column C-YKUJ v v = (v, + 1) . ___ c-Yv Molar fraction of LS in the column cross section of interest

Yv 0

Reflux flow

R v in any cross section of the enriching column

R p v v . E

Vapor flow tiA,,, directly below the feed tray

0

f

Calorific factor of the feed mixture f

fiF2

hS

Ah/,&? YA, w 0

l

r;,

+ v

Ah,,g Enthalpy of the feed mixture F a t the feed and boiling states, respectively Vaporization enthalpy of the feed mixture Molar fraction of LS in GA,,

Vaporflow GA,,in the bottom of the stripping column below the lowest stripping tray, or the packing support grid

yA,, 0

=

Phase equilibrium concentration at the bottom product concentration (approximate)

Refhx flow

kA,,in the

d A , a = GA,a

A

X,

bottom of the stripping column

+A

Bottom product mass flow

* The pressure dependency of the vaporization enthalpies Ai;(,g,Lsand AKl,g,ssare neglected. For equimolar vaporization enthalpies of LS and SS, the vapor and reflux flow rates in both the stripping and enriching columns are approximately constant.

138

2 Distillation and Partial Condensation

If sidestreams Si are withdrawn from the column, their contribution as a heat flux leaving the system, C . &, must be considered in the heat balance (2-49).

si

2.5.2.3 Energy Saving Steps Rectification is an energy intensive thermal separation process. The main task, therefore, in consideration of increasing energy costs and a decreasing supply of conventional energy sources, is to keep the energy consumption of processes as low as possible. This is achieved with respect to produc-

tion capacity, production safety, plant flexibility, and environmental load, with a justifiable expenditure of total cost [2.14, 2.24, 2.451. In general the demands of valid energy saving measures [2.24] are that: 0

0

Energy sources and sinks must be combined in such a way as to minimize the total energy usage With an optimized source/sink combination, the overall energy use matches the tolerated energy losses to the surrounding plus the energy required to maintain the driving temperature gradients of the heat exchanger

Table 2-10. Energy saving steps for rectification.

Choice of optimum operations conditions [2.14]: 0

Operating pressure [2.44]: a reduction in the operating pressure and thereby, a corresponding increase in the relative volatility, leads to a reduced energy requirement and also to heat transfer at a lower temperature level

0

Calorific state of the feed: the feed should be preheated by using available heat in the top and, particularly the bottom products Reflux ratio: a reduction in the reflux ratio leads to a reduced heat requirement but a higher number of separation stages

Choice of the correct column internals [2.14]: A small stage-specific pressure drop is required, particularly with vacuum distillation. Column internals with low pressure drops allow energy savings as well as heat transfer at a lower temperature level. Optimum column combinations [2.141], [2.144]: Selection of the optimum cost separation sequence with direct energy link (see also Table 2-11).

Application of the heat pump principle [2.141], I2.1441: 0

Heat pump with vapor compression: top or bottom product as the working material Heat pump with external auxiliary material flow

Heat transfer at low temperature level: Column operation with reduced pressure, use of low-pressure drop column internals

Application of heat transformation

139

2.5 Countercurrent Distillation (Rectification)

Table 2-10 shows some energy saving measures for rectification processes. In special practical cases, steps may be combined. Energy saving steps often require a higher degree of apparatus “effort” and, with it, higher investment costs. An overall cost estimate is essential for the final plant design. Table 2-11 gives a survey of methods of selecting inexpensive variations of columns arrangements. Figure 2-36 shows simplified arrangements of two columns operated under dif-

ferent pressures. The overhead product of the first column, operated at the higher pressure p l acts as the heating medium of the second column operated at the lower pressure plP With the anticlockwise cyclic process of a heat pump (Fig. 2-37): Polytropic compression 1 2, P A Isobaric condensation 2 + 3 at p N Pressure release 3 4, p N + p A Isobaric evaporation at p A +

0

0

-+

b)

a)

PN

+

Feed

-I

l

r

co1 PI TOP product

TOP

product

+

t

B o t t o m product

1

B o t t o m product

Fig. 2-36. Arrangement of two columns (energy network). a) Serial combination b) Parallel combination CO1 Column with operating pressure p I C 0 2 Column with operating pressure pII(p, >fir) EV Evaporator C Condenses

+

t B o t t o m product

t

140

2 Distillation and Partial Condensation

Table 2-11. Methods for the economic selection (optimum cost) of column arrangements. Structure choice (order of separation, energy network) [2.24]. Heuristic rules [2.401, (2.411 0 Aspire large values of the relative volatility 0 Arrange process such that the top and bottom product remain equimolar as much a possible 0 Separation in order of increasing boiling point 0 Large quantities should be separated early, even if it is a difficult separation 0 Corrosive, toxic and other dangerous substances separated as early as possible 0 Difficult separation to be carried out last 0 Materials which should have a higher purity, should be separated last 0 Desired product should mainly be withdrawn as top product 0 If a cooling media is used for condensation, corresponding separation steps are applied la: Approximate methods (Examples): 0 Approximate method of HARBERT [2.42] 1'' Step: Application of the heuristic rules over equimolar column products and over th process arrangement in order of increasing boiling temperature 2nd Step: The column arrangement with lowest heat requirement is optimal

i=l

Z, Number of separations 0

Approximate method of ROD and MAREK[2.43] With the assumption of mass flows and relative volatility to be approximately constant i the column the column arrangement with the fewest circulating vapor flow is optimal

C VD,,

+

Min!

Analytical methods [2.23], [2.24]: Costing of all possible combinations (separation process sequences without energy network

i

Optimization with regard to minimum total cost

i

Cost estimation of all possible energy networks

1

Fixing the optimal process sequence and energy network

i

Testing the operability

1

Consider alternative separation techniques

1

Consider the heat pump principle

2.5 Countercurrent Distillation (Rectification) a)

QN

141

b)

Effective heat

r-4 3

t

W o s t e heat 5-

QA

Fig. 2-37. Simplified representation of a heat pump process. a) Schematic b) T,s Diagram VC Vapor compressor C Condenser TV Throttle valve EV Evaporator \\\\ Effective heat //// Power consumed I I I I Energy saved

Waste heat QA is used to vaporize a working medium at a low temperature, TA. After polytropic compression of the working medium, during condensation the latent heat Q,, at the higher temperature TN,may be used for heating purposes. If the heat pump process uses the product stream as a working medium; this is known as direct vapor compression. Vaporization and condensation take place in the same apparatus. It is convenient to use the overhead or bottom product as a working medium in a rectification process. Temperature differences between the overhead and bottom products with direct compression (up to ap-

proximately 50 K) are larger than with a heat pump with an auxiliary medium in the cycle (up to 40 K). A temperature gradient between the product and the auxiliary flow is the driving force of the latter cycle and must be taken into account. Figure 2-38 shows a simplified rectification unit including a heat pump; direct vapor compression of flashed bottom product and a heat pump with an external auxiliary medium. Figure 2-39 gives a detailed diagram of a rectification unit with direct vapor compression of the overhead product vapor. Important data for heat pumps for the plant design are listed in Table 2-12.

142 a)

2 Distillation and Partial Condensation

n,

r L

TV

Feed

vc

I

I

SEV -

t

Bottom product

B o t t o m product

Fig. 2-38. Rectification unit with heat pump. a) Direct vapor compression using the bottom product b) Heat pump with an external auxiliary medium CO Column EV Evaporator C Condenser VC Vapor compressor TV Throttle valve SEV Start-up evaporator

The difference between the top and bottom temperature is sometimes large, for example, in vacuum rectification with a high number of separation stages. The economical inclusion of heat pumps is not possible in this case but, if necessary, energy may be saved using a heat transformer. A heat transformer, [2.51-2.53 and 2.1421, is based on the principle of an anti-

clockwise absorption refrigeration machine process cycle (Fig. 2-41). Waste heat at a low temperature T, is raised, with little effort or driving work, to a higher level, where it is used as effective heat. A binary mixture consisting of a very low-boiling and a higher-boiling component, for example ammonia/water, serves as the circulated working medium.

2.5 Countercurrent Distillation (Rectification)

143

a)

I

720 k g l h

9300kg/h

THLi 20 12000 kg/h

IPA 3 0 0 0 kg/h

Fig. 2-39. Rectification unit with top product vapor recompression. Representation according to Sulzer AG [2.48]. a) Process flow diagram b) Example application Separation of an isopropanol/water mixture by azeotropic rectification using toluolene, reduction of energy costs to 70% of these compared to conventional operation by means of a heat pump IPA Isopropanol P Power 0 Heat flow

144

2 Distillation and Partial Condensation

Table 2-12. Important concepts for planning, correlations and data for a heat pump in combination with a rectification unit [2.24, 2.48, 2.1411. Application assumptions and criteria 0 No cheap waste heat available for vaporization 0 Not too large vapor exhaust flows 0 Necessary temperature rise TN- TA < 50 K with direct vapor compression and <40 K with a heat pump and external auxiliary material circulation 0 Heat number Eff > 5 with direct vapor compression 0 Vaporization heat performance 2 2 MW (T, , TN Vapor temperature before and after vapor compression) Application profitability The increased expenditure due to the heat pump must be smaller than the operating cost savings increased investment AZ Cost saving - operating costs of the heat pump > repayment time At I

Eeff

>

[w . AZ

+ cell

1

= Eeff,min

(Gel, CQ specific costs of electrical and steam energy)

Performance parameters Eeff

=

saved or generated effective heat power consumption of compressor

Eeff

=

QN =

Heat number Carnot process equivalent

E,

=

0

Heat pump effectiveness

qHp =

0

Compressor effectiveness

qCom= 0.7.. .0.8

0

0

Heat number (power number)

8 . . , l o (for a rectification plant with heat Neffpump) TN

- TA

7

~

TN

= 0.5 , . .0.7

145

2.5 Countercurrent Distillation (Rectification) Table 2-12. (continued)

Useful heat Q,

required power Neffand pressure rise Apcom over the compressor

0

Effective heat QN = D . Afil,g(A&,,gvaporization enthalpy at T,)

0

Required power

0, MD Mass flow and molar masses of the vapor to be compressed, n polytropic exponent 0

Pressures before ( p A )and after ( p N )the compressor P N = PA

0

Pressure rise Apcom over the compressor, pressure ratio &Corn

=

Apco, APT ApRs 0

+ @Corn A P C ~+IAPCS+ APT + APRS Column pressure drop due to the selected column internals Vapor pressure difference (top product, bottom product boL.. at the same temperature) Pressure difference corresponding to the temperature difference T N - Tn Pressure drop in the pipe work and apparatus (neglibible with a good design)

Pressure ratio (in the validation range of the Trouton rule) (Fig. 2-40)

TT. MR PA PA TS T s . ( T S + W , ) pM Where the variables in Fig. 2.40 are: ps Pressure in bottom of column pT Pressure in top of column ( = p A ) P ATs Boiling temperature difference of the key components at surrounding pressure PS Tn Bottom temperature T, Boiling temperature of the top product pr referred to surrounding pressure ATR Temperature difference over the vaporizer To- Boiling temperature at surrounding pressure Ah,,g = 10.5 . R . To (Trouton rule) In

PN

-=

In

PS

-

+ 10.5 A TS + 10.5 ~

I

A

A

T*

T-

T.

TM

(continued next page)

146

2 Distillation and Partial Condensation

Table 2-12. (continued) Compressor types for vapor compression* Compressor type

Max. pressure (bar)

Max. no. of stages

Max. intake volume (m3/s)

Shell material

Turbo blower

0.1-10

2

20

Radial compressor in pot design one side supported Radial compressor with horizontal, devided housing Axial compressor

ca. 30

2

ca. 10

all weldable materials e. g . , stainless steel, monel, titanium, nickel, etc. (also plated or rubber-coated) Grey cast iron Pigiron

ca. 80

ca. 8

ca. 60

ca. 60

ca. 20

several 100

Grey cast iron, pigiron, cast steel (large types also welded) grey cast iron, pigiron, cast steel (large types also welded)

Example: Comparison of variables for the conversion of a conventional rectification plant for the breakdown of 30 t/h of ethylbenzol/styrene with 50% ethylbenzol into virtually pure components to a plant with direct vapor compression of the top vapor product. * Comparison of variables for column with vapor compression Column top pressure (mbar) Top temperature ("C) Bottom pressure (mbar) Bottom temperature ("C) Reflux ratio Number of theoretical stages AT Vaporizatiordcondensation ("C) Compression ratio Compressor power (kW) Costs for conversion (To) Energy cost savings in comparison to normal operation (070)

*

Representation according to Sulzer AG.

Tray column 67 58 310 106 8.25 54

Conversion of the column to damped packing 67 58 197 93 8.25 54

Conversion of column to Sulzer mellapak 67 58 127 81 6.5 54

15 9.9 4200

15 6.7 3400

15 4.5 2100

100

70

60

30

40

60

2.5 Countercurrent Distillation (Rectification)

In Fig. 2-41, ammonia is expulsed in the stripper (ST) which is heated by some of the waste heat. At a low temperature it is precipitated in the condenser (WKT) and transferred to the evaporator (VT) by the ammonia pump (PIT). The ammonia is vaporized using the waste heat and passed to the absorber (AT). Here it is absorbed by the ammonia weak solution from the stripper, which is cooled in the ammonia preheater (WIT) and the solvent cooler (WT). The ammonia evaporator and absorber are operated at a higher pressure than the stripper and condenser. The heat from the ammonia absorption stage is now at a higher temperature and may be used as effective heat, for example to generate process steam. The heat transformer and the rectification unit are linked together (Fig. 2-42, simplified); the overhead product vapor contains the heat necessary to operate the stripper (ST) and the evaporator (AEV). In the absorber (AT) of the rectification unit, the required heat Q, is generated during absorption of the low-boiling component of the working medium. Since the effective heat of the heat transformer is at most approximately 50% of the heat of condensation of the head product vapor, additional heat is added in the evaporator (EV). Apparatus for a rectification process with less susceptibility to disturbance, pertaining to save energy is outlined in Fig. 2-43. The column is heated with high pressure steam, and its condensate is used after pressure release, at a lower pressure, as a cooling medium for the condenser, where usable low pressure steam is generated [2.54]. For the final evaluation of the economic utilization of energy by means of energy saving steps, an analysis of energetic and exergetic processes is carried out (see Chapter 1.3 and [2.55]-[2.57]). As a result of the energy analysis, energy and enthalpy flowcharts, respectively, give information about 0 The energy introduced by the inflow flow rate, heating medium, etc

0

147

The heats of the products and the cooling medium, the steam generated by waste heat and also heat losses to the surroundings

If there are energy losses the flow diagram is unable to show at which part(s) of the unit the losses occur and how much effective heat is lost. However, in the exergyhergy flow diagram, the effective heat quantity and the location are detected directly as differences between supplied and removed exergy flows.

2.5.2.4 Determination of the Number of Separation Stages and Column Height for Heat and Mass lhnsfer The height Z of a rectification column for mass and heat transfer is given by

if trays as internals are chosen and installed at a tray spacing of Az. N,is the required number of theoretical separation stages, Np is the number of practical separation stages, and Egm is the mean enrichment ratio (mean plate efficiency) of the column. If column internals for continuous phase contact are used (filling and packing material), Z is given by

Z = Nt * HETS

Nt

=-

nt

(2-53)

where HETS is the height of packing equivalent to a theoretical separation stage, n,, (NTS), is the number of theoretical separation stages equivalent to the height of packing. The number of theoretical separation stages Nt required for a rectification problem can be determined by

148

2 Distillation and Partial Condensation

t

Process steam

Ammonic (vapour)

m

Ammonia

Exhaust steam condens

T

3 :.:.:.:.: ;.:.:.:.:.

.....

I

WKT

Exhaust steam I65 100OC)

Ammonia ,(liquid)

cooling water

2.5 Countercurrent Distillation (Rectification)

149

Graphical methods (McCabe-Thiele and where y and x are the molar fractions of the Ponchon-Savarit methods) lower-boiling component 1 in the vapor and 0 Short Cut methods (for example, FENSKE, in the reflux at any cross section of the enUNDERWOOD, and GILLILAND) richment column where both phases are present. (The amount of vapor flow and re0 Step-to-step calculation by computer flux is constant. This is due to the assumption that the molar vaporization enthalpies of both mixture components are equal and neglecting the change of enthalpy of the vaMcCabe-Thiele Method por and reflux. At the top, the adjusted reflux ratio v is constant.) If both mixture components 1 and 2 have The balance line of the enrichment colequal molar vaporization enthalpies and umn (the enrichment line) is described by neglecting the change of enthalpy of the va- its intersection at point A with the 45" diagpor flow and reflux of the column, as onal ( y = x = x E ) and with a given reflux shown in Chapter 1.9.2, the equation of the ratio v at point B on the ordinate (x= 0, balance line for the enrichment section is. y = x,/(v + 1)) in the McCabe-Thiele diagram (Fig. 2-44). If overhead product is withdrawn from V XE the column the slope of the enrichment line (2-54) y=ax+----is always c 1. v+l v+l 0

Fig. 2-41. Single stage heat transformer. Representation according to Krupp Industrietechnik, Grevenbroich [2.53]. AT Absorber WT Preheater ST Stripper WKT Condenser PIT Absorbate pump WIT Preheater VT Absorbate evaporator PLT Dilute solution pump

Characteristic variables, example data Temperature level waste heat: TA > + 40 K Tu Surrounding or coolant temperature Temperature level of useful heat: T N > T A + 30 K Heat ratio: E = QN/QA QN, QA Effective and waste heat Maximum heat ratio E,, (no losses case) Emax =

TA-

Tu

TN-

Tu

TN

*

~

- ., TA

E =

1 . E,,,

, (1exergetic effectiveness)

- TA 5 1.25 , - Tu Example: flA = 80°C; flu = 15°C; flN = 100.. .130°C; E = 0.42.. .0.25; q

Temperature difference ratio:

~

TN

TA

=

0.52.. .0.4

150

2 Distillation and Partial Condensation

C

Fig. 2-42. Rectification unit with heat transformer. Representation according to Krupp Industrietechnik [2.53]. CO Column EV Auxiliary evaporator AT Absorber A P Absorbate preheater ABP Absorbate pump ST Stripper C Condenser ASP Absorbent pump APH Absorbent preheater AEV Absorbent evaporator SP Solution pump ninical examde data. Temperature level, waste heat: Temperature level, effective heat: Condenser temperature: Effective heat/waste heat ratio : Required electrical energy: Cooling water consumption: Savings of 2.7 bar (130°C) vapor: Working media: ,1

80°C, 5.5 MW 120°C, 1.65 ... 2.0 ... 2.14 MW 17.. .27 "C 0.3.. .0.365...0.39 61 kW (pumps), 32 kW (blower) 7.1 m3/h 24800 t/a (7500 h/a) amrnonia/water

2.5 Countercurrent Distillation (Rectification)

151

For the stripping zone as given in Fig. 2-45, the mass balance gives the equation of the balance line (stripping line)

R* = G* + A R*

(2-55)

. x = G * .y + A ' X ,

(2-56)

and hence y=-

R*

A

R* - A

R* - A

*x,

(2-57)

where R* is the reflux, G* is the vapor flow and A is the bottom product flow. Introducing a reflux ratio v* for the stripping zone, y*=-

t

Bottom product

Fig. 2-43. Rectification in energy saving arrangement, use of the heating steam condensate after throttling as a cooling media in the condenser. Production of valuable low pressure steam. Representation according to [2.54]. CO Column R Evaporator C Condenser TV Throttle valve

t r

XE v+l

1.

R* A

(2-58)

it follows from Eq. (2-57) that for the stripping line

With withdrawal of bottom product the slope is always > 1. The first point C on the stripping line occurs at the intersection of a vertical at x, and the 45" diagonal ( y = x = x,). Its position, and the intersection at point S with the enrichment line, are

Fig. 2-44. McCabe-Thiele diagram for a binary system, column with stripping and enriching section. EC Phase equilibrium curve OLE Operating line, enriching section OLS Operating line, stripping section IL Intersection line Mole fraction of the low-boiling compoy nent in the vapor phase Mole fraction of the low-boiling compox nent in the liquid phase

152

2 Distillation and Partial Condensation

must be condensed to preheat it to its boiling conditions, and A & = G* - G = F . ( f - 1)

(2-61)

Hence, the reflux R* is therefore larger than the corresponding sum l? + l? A mass balance over the feed cross section (Fig. 2-46) gives

(R* - ri) . x - (G* - G) ' y - F . x,=

0

(2-62) Fig. 2-45. Material balance of the stripping section. BA Material balance area

governed by the condition of the feed F (concentration x, and thermal state) and the reflux ratio v. The intersection point S gives the conditions of the mixture at the feed cross section (see Fig. 2-44); this is the cross section where the stripping column changes into the enrichment column and vice versa (Fig. 2-46). The factor f is introduced to describe the thermal state of the feed mixture. f.F=R*-R

(2-60)

If the feed is at its boiling temperature, f = 1. If the feed is subcooled, some of the vapor leaving the stripping column AG

where x and y are the molar fractions of the key component (the lower-boiling component) in the liquid and in the vapor, well above the feed cross section. Combining Eqs. (2-60) to (2-62) gives a straight line for the locus of intersections. (2-63) This intersects with the 45" diagonal at point D ( y = x = xF), with the abscissa at point E ( y = 0, x = x F / f )and with the enrichment and the stripping line at point S (see Fig. 2-44). (The coordinates of S, xs and ys describe the composition of reflux and vapor above the feed cross section.) The caloric factor f is obtained from a heat balance over the feed cross section. Since for the preheating of a subcooled feed to boiling conditions the condensation of vapor AG, is necessary.

F . (f - 1)

*

Ah!,g = F * (&s - &F)

(2-64)

From this it follows that (2-65)

Fig. 2-46. Balance of a feed cross section.

where hs and &F are the enthalpies of the mixture at boiling and feed conditions, respectively, and is the vaporization en-

2.5 Countercurrent Distillation (Rectification)

thalpy referred to the conditions at the feed cross section. Different thermal feed conditions and the corresponding position of the locus of intersections are listed in Table 2-13. The point S is fixed by the intersection of the enrichment line and the locus of intersections line, and thus the second point of the stripping line is determined. This is the connection for C and S in the McCabeThiele diagram (Fig. 2-44). The required number of theoretical separation stages Nt is obtained by the construction of steps between the equilibrium curve and the operating lines of the stripping and enrichment zones, as described in Chapter 1.9.2. The example in Fig. 2-44 contains = 4 theoretical steps, including the reboiler in the stripping section, and 2 4 theoretical steps in the enrichment zone. In this case, a feed of composition x, is separated into an overhead product of composition x, and a bottom product of composition x,, of the lower-boiling key component. The stage with a state point on the

153

equilibrium curve closest to S* is said to be the feed stage. If a side stream of composition x, is withdrawn out of the enrichment column, the reflux ratio v reduces to a reflux ratio v,, well below the output . . R-S v, = (2-66) E

s

~

The enrichment line then shows a bend at

x,. Below x, the slope changes according to v, < v resulting in additional number of

separation stages required compared to a separation process without the side stream (Fig. 2-48). If the molar vaporization enthalpies of both of the components of the mixture are not equal, or if considerable enthalpy changes in the liquid fluxes appear, the fluxes are no longer constant in the column section of interest. The stripping and enrichment lines become curves that are determined from point to point in the McCabeThiele diagram. According to HAUSENand

Table 2-13. Thermodynamic state of feed mixture and the resulting position of the intersection line in the McCabe-Thiele diagram. Thermal state of the feed

Temperature

Enthalpy

Calorific factor f

Slope Path of the (see Fig. 2-47) intersection

Subcooled liquid

I9F<<S*

jiF
>1

>O

Boiling liquid

8 F = 8s

hF= h,

1

Wet vapor (partially vaporized) (liquid + vapor)

8 F = !Ys

ii, > r;,

<1

Saturated vapor

8F = 8,

/?F = &s

+ At?,,g

Superheated vapor

19F

>

+ Ah,,,

> 19S

* 8, Boiling point of the feed.

&F

0
(x1

Inclined to the right Parallel to ordinate


Inclined to the left

0

Parallel to abscissa

>O

In c1in ed to the right

154

2 Distillation and Partial Condensation b)

a)

t XA

A

XF X - - c

Fig. 2-47. Intersection line for different feed conditions. OLE Operating line, enriching section SG Intersection lines gradients ( f < 1, f = 1, f > 1, f = 0, f < 0) given by Table 2-13. OLS Operating lines, stripping section

9 f

XF

XS

XE

X - - c

Fig. 2-48. Rectification column with side stream in the enriching- section. a) Schematic representation b) McCabe-Thiele diagram OLE1 Operating line, enriching section above sidestream take off Gradient tan H, = vo/(vo+ 1) OLE2 Operating line, enriching section below sidestream take off gradient tan x2 < tan x1 OLS Operating line, stripping section IL Intersection line EC Equilibrium curve

KIRSCHBAUMand BILLETthe enrichment line is, neglecting the superheating and mixing enthalpy, (linear bubble point and dew point lines in the h , x diagram of MERKEL cording to BITTERL2.59, 2.601, for a variable reflux ratio of the enrichment column and PONCHON, see Fig. 2-11)

Y=

($T 1_

-

1

vo+l

2). x ~-

c

XE

XE

1 --

+

.

vo

vo+l

xE ~-

c

(2-67)

-

A4g2

Ah,,,, - Ah;,,,

c

X 1 -~

C

With chosen and adapted reflux con&tions according to Eq. (2-69), the enrichment line may be calculated by Eq. (2-54). The ordinate intersection of the enrichment line according to Eq. (2-67), is

(2-68)

Assuming linear bubble and dew lines in the Merkel-Ponchon diagram (Fig. 2-11), ac-

(2-69)

* VO

zp

.

where v, is the adjusted reflux ratio at the column head. c is a factor calculated with the molar vaporization enthalPies Al;f,gi and Ahl,g2,where

c=

V

XE

Yo =

v,. (1

-$)+ 1

(2-70)

2.5 Countercurrent Distillation (Rectification)

The following equations are used to describe the stripping line

where x , and ys are the coordinates of the intersection point S of the stripping and enrichment lines. u is given by

where the reflux ratio vs at the feed cross section is (2-73)

155

and MERKEL[2.62, 2.631 is used in its exact form for the mixture being separated. In contrast to the McCabe-Thiele method, the molar vaporization enthalpies of the components are not assumed to be equal and the enthalpy of mixing not neglected. The Ponchon-Savarit method, therefore, exactly determines the number of theoretical separation stages by taking into account the real caloric conditions of the separation column. However, it is more costly and more complicated to handle than the McCabe-Thiele method and the PonchonMerkel diagram is only readily available for use in a few systems. CHON

Short Cut Method of FENSKE, and GILLILAND UNDERWOOD This method is based on the assumption of constant flow rates, R and G, constant relative volatility a1,2and an infinite reflux ratio v. According to FENSKE and UNDERWOOD [2.64] the required minimum number of theoretical separation stages Nt,min for a binary rectification is

A comprehensive representation of different approaches to obtain the operating lines of rectification columns is given by BITTER in [2.59, 2.601. In practical column operation at higher loading, liquid is entrained with the upflowing vapor. The ratio of the entrainment and the amount of upflowing vapor influences the local reflux ratio and thus the position The reboiler is not taken into account. An infinite reflux ratio v = 00 means no of the operating lines and the enrichment product is withdrawn at the head of the colratio. umn (E = 0); all condensed overhead vapor is returned to the column. In the McCabeThiele diagram, stripping and enrichment Ponchon-Savarit Method lines coincide with the 45" diagonal and therefore the concentration of the feed x, The method of PONCHON and SAVARIT [2.61] is a graphical determination of the number has no influence on the minimum number of theoretical separation stages in counter- of stages Nt,min. For similar values of the relative volatilflow columns. It is particularly applicable to binary mixtures in rectification pro- ity at the top and bottom of the column, the geometric mean is cesses. Exact results are obtained when the a1,2,Eand a1,2,A, enthalpy-concentration diagram of PON- used in Eq. (2-74).

156

2 Distillation and Partial Condensation 1.0

0.6 0.L

t

0.2

0.1

0 0.0; 0.0

Fig. 2-49. Gilliland's diagram for theoretical stages and reflux requirement. EOR Economic operating range

a1,2 = h

,

E

. QV,A

(2-75)

Figure 2-49 shows an empirical relationship between the number of theoretical separation stages N, and the reflux ratio v introduced by GILLILAND [2.64]. For a given rectification problem the minimum reflux ratio vmin and the minimum number of separation stages Nt,,,in may be determined by the McCabeThiele method or calculated by Eqs. (2-74) and (2-77). With these, the required number of theoretical stages for any chosen reflux ratio v is found from the Gilliland diagram. The best economic operating range of rectification columns is

The described short cut methods may also be applied for difficult rectification problems, for example, the separation of azeotropic mixtures. A suitable transformation of coordinates for the equilibrium

Fig. 2-50. Transformation of coordinates, azeotropic binary systems. x, y Coordinates of the azeotrope mixture <, Coordinates of the nonazeotropic pseudomixture with OT as the coordinate origin Example: ethanol/benzol [2.65] 1 > a > 0.240 for a = y . (1 - x)/[x.(1 - y ) ] ; a = 0.252 k 0.012 for a = r] . (1 - O / [ < . (1- r ] ) ] .

composition of the azeotropic mixture is required, thus the relative volatility of the transformed system should be almost constant in the range of interest (Fig. 2-50).

Calculation of the Number of Stages, Concentration and Temperature Profiles If a multicomponent mixture with pronounced nonideal behavior is to be separated, graphical and short cut methods fail. They also fail if a key component calculation gives insufficiently accurate results. The number of stages or the height for heat and mass transfer of a rectification column are determined by the aid of a computer. Computations are based on an equation system (see Chapter 1.9) consisting of 0 0 0 0

Mass balance equations Phase equilibria relationships Enthalpy balance equations Stoichiometric conditions of the sum of concentrations

2.5 Countercurrent Distillation (Rectification)

157

For equilibrium curves including a turning point, vmin may be found in such a way that the balance line is tangential to the equilibrium curve (Fig. 2-51). Since the number of separation stages cannot be infinite the real reflux ratio must be v > vmin. The operating costs C , of the rectification unit (energy costs for heating and cooling) increase with increasing reflux 2.5.2.5 Minimum Reflux Ratio, ratio. The investment costs C, of the colOptimal Economic Reflux Ratio umn increase with increasing number of installed separation stages and decrease with The reflux ratio v, where increasing reflux ratio. The investment costs R increase slightly with large reflux ratios due v=(2-76) to the high manufacturing costs of columns E with large diameters. The economic optimum reflux ratio vOpt, is the most important operating variable in a rectification process with a set feed condi- is given by the requirement of minimal total tion and operating pressure. v may be varied costs C, = C, + Co of the rectification between the minimum reflux ratio vmin and unit, and is graphically determined as shown in Fig. 2-52. Several variations in the 00. The minimum reflux ratio vminis the reflux ratio requiring an infinite number of rectification unit with different reflux raseparation stages in the stripping and en- tios, are calculated obtaining the curves C,(v) and C,(v) in Fig. 2-52. The richment column to realize the separation C,(v), economic optimum reflux ratio vOpt is thus desired. At constant relative volatility a1,2 found at the minimum of the total costs and with a feed at boiling conditions, vmin curve, C,(V). is according to FENSKE and UNDERWOOD, In practice, operating reflux ratios of v = (1.05 2) vmin and sometimes higher are used. The system is nonlinear and must be solved using a reasonable estimate of the iteration variables (for example, the condition of the product flows at stage to stage calculations, mass, concentration or temperature profiles) [2.66-2.681.

2.5.2.6 Feed Stage The minimum reflux ratio may also be determined graphically from the McCabeThiele diagram, from the requirement that the stripping and enrichment lines intersect at point S* (Fig. 2-44). vmin results from the ordinate intersection yo,min with the enrichment line, giving

v min . =-- xE

1

(2-78)

Yo,min

With unequal molar vaporization enthalpies, vmin is calculated from Eq. (2-70).

The location of the feed stage is chosen in such a way that as few separation stages as possible are needed in the stripping and enrichment column. Figure 2-53 shows examples of early, delayed and correct feed locations in the sense of a minimum number of separation stages. An early feed location increases the number of separation stages required in the enrichment column, whilst a delayed feed location increases the number of separation stages in the stripping column. In the given example, one separation stage is saved when the correct feed location is chosen.

158

2 Distillation and Partial Condensation b)

a)

XA

XF X+

Fig. 2-51. Graphical determination of the minimum reflux ratio. a) Intersection point method b) Tangent methold OLE Operating line, enriching section OLE, Tangential operating line, enriching section, from A out through the equilibrium curve OLS Operating line, stripping section IL Intersection line EC Equilibrium curve T Tangent contact point

2.5.3 Discontinuous Adiabatic Rectification

Fig. 2-52. Reflux ratio for optimum design. CT Total Costs C, Investment costs C, Operation costs

In discontinuous rectification a liquid mixture is charged to a still (Fig. 2-54) and after heating the mixture to boiling point, rectification occurs in an enrichment column operated on top of the still. The vapor leaving the enrichment column is then condensed in a condenser. Some of it is passed back as reflux to the column, the rest is continuously withdrawn as product. After the rectification is finished, the bottom product (distillation residue) is removed from the still. Discontinuous rectification is less costly with respect to equipment than the continuous variation and multicomponent mixtures may even be separated in a single column. Due to the considerable dead times for charging, discharging and heating of the

2.5 Countercurrent Distillation (Rectification)

XA

XE -X

Fig. 2-53. Influence of delayed and early feed mixture entry on the required number of separation stages. a) Early feed b) Delayed feed c) Correctly chosen feed Z Feed stage

RC ~-

Q

159

Fig. S C TC TR RC

2-54. Batch distillation unit. Still Condenser Top product cooler Top product receiver Rectification column

160

2 Distillation and Partial Condensation

still, and the tendency to use continuous operation in chemical plants, discontinuous rectification is seldom used. It is mainly employed for small amounts of mixture, which change irregularly with time and mixture composition. In practice, there are two variations of discontinuous rectification. In the first variation, the reflux ratio at the head of the column is kept constant. The composition of the top product varies with time and with the composition of the still contents. This variation is preferred for the separation of multicomponent mixtures. At the start of the rectification process, a large reflux ratio is used to remove low-boiling components of the mixture during this “first run”. If the desired product component appears in the overhead product, the reflux ratio is decreased. During the “main run” a low reflux ratio is kept constant. At the end of the main run the reflux ratio has to be increased again to obtain as much as

possible of the remaining product from the still mixture. Figure 2-55 shows a typical curve of the temperature of the top product against the amount of distillate, for discontinuous rectification of a ternary mixture. In the second variation, the reflux ratio is continuously adjusted to maintain a constant overhead product composition. This variation is preferred for fractional separation of binary mixtures. Exact mathematical treatment of discontinuous rectification is not simple due to the changing component concentrations in the still contents and the overhead product, and the concentration profile of the column with time. Therefore, only a few important relationships which apply under certain assumptions are given in the following chapter. More details are found in [2.1-2.31.

tI

During a discontinuous rectification process of time t, a liquid mixture B, with a concentration of x, is separated into a top product Eg of concentration x, and distillation residue B, of concentration x,, (see Fig. 2-54). Molar fractions x,,, xE and x, refer to the low-boiling component as the key component. fg is the separation time, fixed by the order of events of the total process. The important cases for process control are constant vapor load of the column and constant heat supply over a time period. These are discussed in detail in [2.2]. If the column is operated with a constant reflux ratio v, a material balance of the column gives analogous to Chapter 2.2

Top product flow

-

Fig. 2-55. Path of the temperature in the top of a column, which depends on the distillate top product flow, for a three component batch distillation process. Area A, B; C, D; E, F Withdrawal of virtually pure product fractions with low reflux ratio Area B, C ; D, E Withdrawal of products of intermediate composition. The intermediate fractions are usually collected separately and feed to the still together with the following charge.

2.5.3.1 Amount of Overhead Product

(2-79)

and the total amount of overhead product

Eg is

2.5 Countercurrent Distillation (Rectification)

Eg = B, - B,

= B,

*

Eg The simplest ways of solving the integral expressions in Eqs. (2-79) and (2-80) are graphical or numerical means. Table 2-14 shows the order of the steps which are required. The time mean composition of the distillate x ~is , ~

fiE

the top of the column and concentration x,, total amount of withdrawn top product during the rectification process molar enthalpy of the top product

By neglecting the liquid contents in the column and condenser, according to Eq. (2-82), the amount of top product E ( x ) is xB, -

(2-81) If the reflux ratio v is adjusted at a time to obtain a constant overhead product concentration X E , , the total amount of produced overhead product is

161

E=B,*-

xEa -

(2-84)

Table 2-15 describes the course of action to evaluate Eq. (2-83). If the reflux ratio v is kept constant at the column top, the total heat requirement Q of the column is

(2-85)

2.5.3.2 Heat Requirement With constant top product concentration

xEa,the total heat requirement Q is

Now A&,,g and f i E depend on X , and hence on E. Analogous to Eq. (2-84), the amount of top product E ( x ) is (2-84)

where Ah,,, enthalpy of vaporization of the top product, referred to the pressure at

x

~is the , ~time mean molar fraction of low-boiling component in the top product, which is analogous to Eq. (2-81)

Table 2-14. Step-by-step evaluation of the integral terms in Eqs. (2-79, 2-80). 0 0 0 0

I

Determination of the number of theoretical stages Nt (see Chapter 2.5.3.4). Selection of x values in the range of interest (xe, > x > xe,) Graphical determination of x, corresponding to x, considering v = const. and at constant Nt Plotting of xE(x) and evaluating the integral I

162

2 Distillation and Partial Condensation

Table 2-15. Step-by-step determination of the heat requirement Q with constant top product composition x,. 0 0

0

Selection of the reflux ratio v in the interval v, < v < v, where v, and v, are the initial and final values of the reflux ratio Graphical determination of the still concentration x corresponding to v using the McCabeThiele diagram at fixed number of separation stages Nf(see Chapter 2.5.3.4) Calculation of E using Eq. (2-84) and plotting v ( E ) + 1 against E Graphical evaluation of the integral term in Eq. (2-83) and calculation of Q

xE,m =

B;x

Bil

-B*x

B,-B

(2-87)

According to Eq. (2-79) the still contents B is given by

s-)

XB.

B=B,.exp(-

dx

x, - x

(2-88)

The course of action to evaluate Eq. (2-85) is described in Table (2-16). In discontinuous rectification practice, for the important cases x,, = const., v = const., the heat requirement calculated from Eq. (2-83) or (2-85), is approximately the heat supplied to the distillation still during the total time period of rectification tg. The change in the heat required during the rectification process, resulting from the change of the liquid contents in the still from B, to B,, is neglected. Therefore, the calculation of the heat exchange area and

the flow of the heating medium already includes a design safety factor. To heat the still contents after filling B, to boiling point, the heat requirement Q,, is

where ha and hs are the molar enthalpies of the mixture B,, to be separated at the filling temperature and boiling point, respectively.

2.5.3.3 Still Diameter, Free Vapor Space, Column Diameter

The diameter dof the distillation still and the height of the free vapor space Z , may be calculated by means of Eqs. (2-7)-(2-10) given in Chapter 2.2. The calculation of the column diameter is discussed in Chapter 2.5.5.

Table 2-16. Steps in the determination of the heat requirement Q with constant reflux ratio. 0

0

Determination of B for Eq. (2-88) as described in Calculation of x ~using , Evaluation of the integral tion of Q

assumed x values within the interval xB, > x > x B , using Table 2-14. Eq. ~ (2-87) and E using Eq. (2-86) term in Eq. (2-85) with fixed A&,,g( E ) and h E ( E )and calcula-

163

2.5 Countercurrent Distillation (Rectification)

2.5.3.4 McCabe-Thiele Method to Determine the Number of Theoretical Separation Stages With discontinuous rectification processes, vapor generated in the distillation still is guided through the column. The column is operated in enrichment mode, the reflux is part of the condensate from the condenser and flows countercurrently with the upflowing vapor (Fig. 2-54). The vapor concentration y and reflux concentration x at a particular cross section, are linked by the equation for the enrichment line (Eqs. (2-54) and (2-67)). The required number of theoretical separation stages is derived by constructing stages between enrichment line and equilibrium curve, as shown in Fig. 2-56 for a binary mixture. It has to be kept in mind, that due to the continuous withdrawal of top product, the composition of the still contents also continuously changes the mole fraction, the low-boiling fraction decreases from the initial value x,, of the rectification process to the value x at the point in time of interest. The mole fraction of the low-boiling component in the

top product decreases from x,, to x, during the rectification time t with a constant number of separation stages and constant reflux ratio v. Otherwise, to produce a top product of constant concentration x,,, with a constant number of separation stages, the reflux ratio at the top of the column has to be increased and adjusted constantly. The minimum reflux ratio vmin for a mole fraction X , of the low-boiling component in the still, is (2-90) where X, and y B determine the coordinates of the intersection of the enrichment line and the equilibrium curve.

2.5.4 Semicontinuous Adiabatic Rectification In semicontinuous rectification, a fractionally generated vapor mixture is separated into continuously downflowing fractions in

b)

a1

t X - - c

Fig. 2-56. McCabe-Thiele-diagram for calculating the number of theoretical stages for the batch distillation of a binary mixture. a) Constant reflux ratio v = const b) Constant composition of the overhead product xE, = const

164

2 Distillation and Partial Condensation

a rectification column consisting of enrichment and stripping sections (Fig. 2-57). Since the fraction of the low-boiling component in the still continually decreases during the rectification process, the composition of top and bottom products may be kept constant only by continuously increasing the reflux ratio.

containing ethylene glycol, is charged directly back to the reactor. Mathematical treatment of semicontinuous rectification processes is found in [2.2].

2.5.5 Determination of the Column Diameter The cross-sectional area A Q and diameter d for a rectification unit, result from a flow equation in which the vapor is the reference phase

Fig, 2-57. Semicontinuous distillation unit. DS Distillation still or batch reactor RC Rectification column SC Stripping column C Condenser R Reboiler

Semicontinuous rectification processes are of practical importance for vapor mixtures generated in a discontinuous chemical reaction. These then have to be separated into a component continuously fed back to the reactor and a discharging fraction. For example, the main task during the transesterification of dimethyl teraphthalate with ethylene glycol is to continuously separate methanol from the reaction mixture. The reaction equilibrium is then shifted toward the product. Methanol is the top product of the column which is connected directly to the transesterification reactor. The high-boiling, or heavy, bottom product,

where V,,,,, is the maximum possible effective volumetric flow rate of vapor through the column, w ~is the , maximum ~ ~ ~ allowed velocity, referred to the free crosssectional area of the column, Gmax is the maximum possible flow rate, Q, is the vapor density, and Mg its molar mass. With Gmaxin kmol/h, w ~ in ,m/s,~ operating ~ ~ temperature T in K and operating pressure p in bar, the column diameter d is

d

= 0.00542 *

(2-92)

Ideal behavior of the vapor phase is assumed. It is recommended to compare the result from the calculations for four reference cross sections; the cross section at the column head, the cross sections above and below the feed tray and the cross section of the column bottom. The maximum allowed vapor velocity w ~depends , ~ on~the~type and geometry of the column internals, on the reflux load, and the properties of the phases in contact.

2.5 Countercurrent Distillation (Rectification)

2.5.6 Internals in Rectification Columns In order to reach the most possible intensive contact between the counterflow phases (vapor and reflux) and hence good mass and heat transfer, the rectification column is equipped with internals. Internals include trays, rotating devices and packing. Columns with stepwise phase contact, e. g., tray columns, and columns with rotating internals, and wetting columns with a continuous phase contact, e.g., packed columns can be distinguished. In tray columns with controlled flow, the liquid reflux passes via an inlet downcomer from the upper tray to the horizontal tray immediately below it. The liquid then flows across the tray as a continuous phase and flows from this tray to the next lower tray via a downcomer (Fig. 2-58a). The desired liquid content is ensured by installing a weir in front of the downcomer. Vapor enters the tray directly through openings in the bottom of the tray (drilled holes, slits, throats, etc.) or via rigid or flexible caps where the vapor forms bubbles which are dispersed in the liquid. In overpressure operation, with large weir heights and small vapor loadings, a bubbling bed forms on the trays. In vacuum operation, with small weir heights and large vapor loadings, the reflux liquid is dispersed in droplets and a droplet bed forms. With medium loading conditions, depending on the vapor load a two phase region forms between the trays, consisting of bubbling zone and a spray zone in which liquid droplets are entrained (splash zone). Heat and mass transfer take place in both bubbling and spray zones. Therefore, the transfer area is the total surface of the vapor bubbles and the liquid droplets. On trays with controlled vapor and liquid flow (reflux), the basic flow pattern is cross flow. In columns without controlled liquid guidance, the flow pattern is counterflow.

165

No phase contact between the spray layer and the next upper tray is possible in columns with forced flow control; both phases flow separately in counterflow. The maximum vapor loading of a tray column has to be adapted to the liquid loading, so that no liquid is carried by the vapor to the next upper tray. With minimum vapor loading, reflux should not drain through the floor drilling and the column should not become empty. In packed columns, liquid reflux flows as a falling film, or as a streamlet, from top to bottom counterflow to the upflowing vapor. Both liquid and vapor phases are in continual contact (Fig. 2-58 b and c). Mass and heat transfer occur at the inside and outside surfaces of the randomly packed filling material or the arranged packing elements in reflux film. The exchange area is the surface area. In the case of spraypack fabrics, the reflux liquid is sprayed. The contact area is the total surface area of the liquid droplets. Maximum gas and liquid throughput rates of packed columns depend on the type of packing and packing geometry, the relative free void fraction and the physical properties of the mixed phases. The column becomes flooded beyond a certain volumetric flux of the vapor at a given flux of the reflux. A controlled counterflow of the phases then does not exist; the separation efficiency of the column is reduced drastically. In rotary columns, the reflux is sprayed by means of a rotor with a funnel shaped distributor or is distributed by means of a rotor as a film on a heated tube wall. The phase contact upflowing vapor-downflowing liquid is stepwise in the first case and continuous in the latter (Fig. 2-58d). In Table 2-17, important criteria for the selection of column internals are listed. Technical considerations include separation efficiency, maximum loading and pressure drop. To date, it has not been possible to

166

2 Distillation and Partial Condensation

tD 1

C)

d)

1 --

RC

RO

4

Vapor phose

Liquid phose

Fig. 2-58. Rectification column internals. a) Tray column D Downcomer ID Inlet downcomer BZ Bubbling zone SZ Spray zone b) Random packed column R P Random packing c) Wetted packing column MP Regular mesh or extended metal packing SP Spraypack d) Rotary column RO Rotor with funnel shaped distributor RC Reflux collector ? Vapor phase 1 Liquid phase

large variety of rectification trays, random and structured packings available. In general, trays should be used as column internals, if: 0

0

0

0

Large throughputs require a large column diameter With throughput variations sufficient flexibility is important Larger tray pressure drops under coarse vacuum, normal pressure, or overpressure operation are tolerated A possibility of contamination and incrustation occurs

Random packings in smaller column diproduce internals with maximum exchange ameters are characterized by efficiency accompanied by maximum loading and minimum pressure drop. Instead, 0 High specific loadability the column internals have to be chosen ac- 0 Good separation efficiency cording t o the particular requirements of 0 Small specific pressure drop the rectification problem. This explains the 0 Low costs

2.5 Countercurrent Distillation (Rectification)

167

Table 2-17. Considerations for the assessment and selection of column internals.

0 0 0 0

0

Effectiveness, separation effect (enrichment ratio, NTS, HETP-value) Loading ability, loading range (highest allowed vapor velocity, smallest possible vapor velocity, reflux stream density) Pressure drop, equivalent pressure drop for a theoretical stage Flexibility for loading variations Sensitivity against fouling and crusting Possibility of manufactoring of materials resistant to different mixtures Acceptable column costs

This type of filling material is particularly useful with normal pressure and under coarse vacuum operation. Structured packings are better than random packings with respect to separation efficiency and loadability. Since they cause smaller specific pressure drops, though at higher costs, they are suitable for difficult separation tasks in vacuum operation. A detailed discussion of column internals and a rating with respect to costs are given by BILLETin [2.2, 2.14, 2.691. Table 2-18 gives a detailed overview of technical columns in counterflow of gasliquid, designed by MANTEUFEL [2.108]. 2.5.6.1 Column Trays Liquid Flow The flow pattern on column trays is usually controlled. Due to the design arrangement of the entry and exit, including an exit weir, the reflux liquid flows across the tray while the vapor flows upward (Fig. 2-59). Vapor and liquid therefore flow in a cross currentcounter-flow manner through the column. With trays without controlled liquid flow, vapor and liquid flow countercurrently through the same openings in the tray. Holdup of the liquid on the tray for sufficient heat and mass transfer is obtained only under certain flow conditions.

Important Types of Column Trays

Column trays are horizontally arranged elements of plate, formed and strengthened by channel sections. In general, trays are assembled as individual plates or as packages of several trays, mounted at the column wall with a liquid-tight seal (Fig. 2-60). Trays may be accessible by means of manholes in the column wall and service openings on the trays for maintenance, cleaning and reconstruction purposes (Construction details are found in [2.2, 2.70, 2.711). The following groups of column trays are important : Trays with drilling or slits in the base plate. Liquid flow controlled across the base plate or drops downward through the base plate openings where the vapor flows counterflow upward, e. g., sieve tray, turbogrid tray) Trays with vapor flow through throats, or chimneys covered by bells, hoods, or caps, dipped in the liquid. The liquid flows in a controlled manner across the base plate (e. g., bubble cap tray, cross flow tray, Thormann tray, Streuber tray) Trays with drilled holes in the base plate, which are covered by movable and adjustable load valves. Liquid flows in a controlled manner across the tray (e. g.,

8. Overview of full-scale technical columns with countercurrent flow of gas and liquid phases *.

Empty

Empty

Bubbles

Drops

Continuous Countercurrent

Continuous Countercurrent

Bubble column

Spray column

Only at top on

with

Tray column without reflux drain

Stagewise Countercurrent cross flow Bubbles/drops

-

Stagewise Countercurrent Bubbles/drops

Horizontal directors (trays) (radial) Liquid weir on each tray

Bubble cap trays Sieve trays (hole trays) Tunnel-cap tray (Thormann) Valve tray Perforated tray Sieve turbogrid tray Sieve bubble cap tray

Only at top

Wire mesh tray (Turbogrid) Ripple tray Perforated tray Kittel tray Corrugated sieve tray

Random packed column

Continuous Countercurrent Film/bubbles Random packing

Structured packing

Continuous Countercurrent Film Vertical dividers (axial)

C C Fi dr C

Sulzer packing P Stedrnan packing Spraypack Braided metal packing Braided metal spiral packing (Montz)

Raschig rings Tube bundle Saddel (Berl-) (Kuhn) Pall rings Tube bundle (Brauer) Intos-Ring Plates Perro-Ring (Adolphi, G) Intalox Grid Novalox Square grid Ralox Spring Interpack Random (Montz) Trickle wire weave packing packing (Stage!

Every 1-3 m O or only at top

Only at top or several intermediate distributors

Only at top or several intermediate distributors

Jet tray Centrifuge tray Perform contact tray MD-sieve tray Film tray Metal

Metal

(i. e., Multifit, Hyperfit)

Shaped-metal trickle packing

100-300

100-200

Metal Ceramic

Ceramic Metal Plastic

60,

450-800

Po (c M Pl

Metallic weave Plastic Porcelain

ing

2-3

2-4

1-2

1-2

2-2.5

1.5-2.5

1-2.5

0.5-2 1-3

3-

10-6

10-2.5

1-

5-2

5-0.5

2-3

2-3

2-

olumn

hange [2.109]

[2.109]

[2.731,[2.74]

[2.73],[2.74]

5

10

0.5- 0.3

0.5-0.25

1-0.3 1.5 i2.1101

0.1-0.4

0.1-0.2

P.21

v.21

1

1

0.

ture : olume-specific surface (m2/m3) loading factor F, = w g . f separation stages: number of stages per unit column height n, = N , / Z olumn volume: per unit vapor volumetric flow and per separation effectiveness required column volume hange: ratio of effective gas velocity through the gas openings and gas velocity related to the column cross sectional ntation according to MANTEUFEL [2.108]

6

170

2 Distillation and Partial Condensation a) Section 1-1

1

Vapor

4

Section 11-11

Reflux

Fig. 2-59. Column cross-flow trays. a) Single-pass tray, available up to ca. 2 m diameter b) Four-pass tray for large column diameters and high liquid loading c) Radial flow tray d) Stepped tray to even out different liquid levels at high liquid loadings e) Reverse flow tray for small liquid loadings t Vapor 1- Reflux

2.5 Countercurrent Distillation (Rectification) 2

L

171

c’

3

Fig. 2-60. Attachment and sealing of trays. Representation according to Montz, Hilden. a) Tray with attachment device b) Tray sealing with spiral spring c) Function of the tray sealing 1 Shell 2 Cage of spiral spring 3 Spiral spring 4 Tray 5 Spacer 6 Bolt 7 Bolt attachment 8 Tray support ring

0

Glitsch valve tray, Stahl valve tray, Koch flexitray) Specially constructed trays with specially designed vapor openings, cause an increase in the cross flow due to a deflection of the vertical vapor jet (e. g., perforated valve tray, Kittel tray, jet tray, perform contact tray)

Well known types of column trays are presented and explained in Tables 2-19 and 2-20 with respect to their operating mode and important design data. ~

Figure 2-62 shows different areas of a cross section of a tray which is common to all cross flow trays. These areas must be considered in the fluid mechanical design of the tray. Table 2-21 compares some important column trays. For the detailed evaluation of economical efficiency or optimization calculations, experimentation with the separation mixture in pilot-scale plants is required [2.69, 2.721.

Loading Range, Operating Range The operating or loading range of a tray column is limited by the type and geometry of the trays and also by the properties of the phases in contact. From the viewpoint of large throughputs and good flexibility, a tray column should have the largest possible operating range. Figure 2-63 shows qualitatively the operating range with controlled flow. The following loading limits must be considered :

172

2 Distillation and Partial Condensation

Tab. 2-19. Commonly used column trays with explanation and brief description of the design and operation [2.73]. ~~

~

~

Tray Explanatory sketch

Design and operation principle

Column trays without guided liquid flow Sieve tray without downcomer

Sieve plate without downcomer, countercurrent vapor and liquid flow through the same openings. Liquid rains in dispersed droplets from tray to tray and sprays when landing on next tray. Minimum vapor velocity must be obeyed to keep a liquid layer on the tray as the heat and mass transfer area.

[2.74]

Ripple tray

Approximately sinusoidally ribbed sieve plates or shaped metal surface. In order to guarantee a good liquid distribution, plates are placed with a 90" shift, with reference to the plate above. Countercurrent vapor and liquid flow, the vapor mainly flows through the ribbed plate peak openings and preferentially the liquid drops trough the plate valley openings.

[2.74]

Grid or shaped plate, countercurrent vapor and liquid flow, both phases flow interchangeably through the slits. The liquid is thereby dispersed as droplets. Trays used with and without liquid guidance.

[2.74] [2.75] [2.76]

References

, 0

0 0 0 0

0 0

1: .

Turbogrid tray

2.5 Countercurrent Distillation (Rectification)

173

Tab. 2-19. (continued) Tray Explanatory sketch

Design and operation principle

Kittel tray without downcomer

Shaped metal trays, vapor and liquid flow through the same slits. The slits are arranged such that the vapor imposes the direction of movement of the liquid. The liquid on each tray is pushed from the edges to the middle and drains there preferentially. On the neighboring trays the flow direction is opposite.

[2.74]

Column trays with guided liquid flow

Sieve plates with tray openings for the vapor flow and feed or drain appliances. Minimum vapor velocity must be obeyed such that liquid does not rain through the trays. As in the case of sieve slit trays, additional slits can be stamped into the surface of the sieve plate, through which some of the vapor is reflected in the direction of the liquid flow. This is advantageous with large liquid loadings. More downcomers can also be ordered as in the case of MD (Multi Downcomer) Trays.

12.21 [2.74] [2.77] [2.78] [2.91]

Sieve tray

ID Inlet downcomer OD Outlet downcomer

References

(continued next page)

174

2 Distillation and Partial Condensation

Tab. 2-19. (continued) Tray Explanatory sketch

Design and operation principle

Bubble cap tray')

Cross-flow tray with openings or chimneys for the vapor flow, covered by caps with smooth, slitted or peaked edges. The upflowing gas phase is directed up through the caps and flows out in parallel to the tray base through the liquid, such that a bubbling layer for heat and mass transfer forms. The cap submergence remains fixed by the distance between the top of the downcomer and the bottom of the cap, and is the submergence of the vapor path through the bubble layer. A decrease in the bubble cap diameter increases the separation effectiveness but enables an increase in the loading. The modern flat or low-riser cap designs show good effectiveness with higher loadings and low pressure drops. For higher vapor loadings, rectangular ( - 100 x 50 rnm) cross-wise arranged caps are used (cross-flow cap trays).

I

I

Section 1-1

1

2

a

I

-b

3

-- a nnnn

L

b C

6 d

e

a

Cap designs 1 Conventional cap') 2 Low-riser') 3 Sigwart cap') 4 Umbrella cap') 5 Varioflex cap') 6 Varioflex cap with valve plates2) (nozzle discharge disc) a Vapor opening (tray opening, chimney) b Cap c Vapor outlet slit with guide plate d Nozzle disc lift stop e Valve plate (nozzle disc)

References P.21 [2.74] [2.80] [2.81]

2.5 Countercurrent Distillation (Rectification)

175

Tab. 2-19. (continued) Tray Explanatory sketch

Design and operation principle

Channel tray3)

Cross-flow trays work to the same principle as bubble cap trays; they have length-wise vapor chimneys which are fixed by the U-shaped channels and hood covers. The channel breadth, chimney breadth, chimney height and hood submergence depend upon the operating conditions. Parallel (STREUBER system) or channels mounted crosswise to the (THORMANN system) liquid-flow direction are used. Annular channels are also used, in which case the flow on the tray is radial. The type and design of the head slits define the direction of the vapor flow and, therefore, an acceleration or deceleration of the liquid ( ~ O R M A N Ntray). 1 Channel tray with cross-liquid flow defined by the channels and hoods (THORMANNsystem) 2 Liquid flow through a channel tray (THORMANN system) 3 Addition of a hood with alternating folded slits 4 Slit sketch of a vapor chimney and hood

References K2.21 [2.85]

(continued next page)

176

2 Distillation and Partial Condensation

Tab. 2-19. (continued) Tray Explanatory sketch

Design and operation principle

References

~

KSG cross-flow tray3) 1

KSG-S

Cross-flow tray with punched in vapor chimney with rounded edges (types KSG-S, KSG-0) or with direct punched-out gas outlets (type KSG-E). Chimneys are covered by hoods with outlet slit. To avoid an impact on opposite vapor fluxes neighboring vapor fluxes cross each other under a 90" angle due to the opening arrangement of the exchange devices (positive effect on pressure drop and entrainment). Depending on the vapor load on the KSG-S tray, vapor flows through either one or both output slits arranged on top of each other (wide operating range; operating characteristic like a valve tray). 1 Cross-flow tray 2 Vapor chimney with hood

KSG-0

-& KSG-E

Perforated valve tray4)

Froth

rl

Froth

Tray with slits arranged at side. Vapor or gas is guided through the horizontally flowing liquid. Thus, with a small liquid holdup a long contact time is achieved. Type A Liquid always flows from the outer edge radially across the tray to the center (radially in the same direction)

[2.90]

2.5 Countercurrent Distillation (Rectification)

177

Tab. 2-19. (continued) Tray Explanatory sketch

Design and operation principle

1

S-Tray (Uniflux tray)

1-

Valve tray5)

bl

References

Type B Alternative liquid flows from outside toward the center and from the center toward the Froth outside (radial opposing)

Cross-flow tray with across-mounted S-shaped elements including trapezoid slits. Liquid accelerated in downcomer weir direction by the vapor leaving the chimneys.

12.21 [2.74]

Cross-flow sieve tray with large holes (hole diameter = 40 mm). The holes are covered by guided movable ballast valves, plate valves or lever valve (float valve). Valve lift and, therefore the opening for the vapor flow is adjusted automatically by the momentum of the upflowing vapor. While exiting the valve vapor is diverted under the valve cap to flow parallel to the tray base plate into the liquid. Thus good vaporliquid mixing is provided by the valve caps over a wide vapor load range. To avoid draining of the tray a minimum vapor load is required. Different valve designs are: 0 “Three-leg” valve (glitsch valve, ballast valve) guided by three internal legs

12.21 [2.74] [2.82] [2.87] [2.88]

(continued next page)

178

2 Distillation and Partial Condensation

Tab. 2-19. (continued) Tray Explanatory sketch

Design and operation principle

CJ

9

L

7

1

0

2

6

5

9

3

L

6

5

L

?

1

0 0

2

6

5

3

6

5

L

Square “four-leg” valve (Speichim capsule tray) Valve disc with smooth or venturi tube shaped vapor entry guided in a cage mounted on the tray (Koch Flexitray, Stahl Varioflex-cap with valve disc) Rectangular float valve (Nutter float valve tray) Cross-wise mounted area valves guided by internal legs (Montz cross-flow valve tray) Valve disc attached to the grid (Krupp “grid valve”) a) Valve tray (schematic) b) Operation principle of a ballast valve in fully open position c) Assembling of a ballast valve

1 Ballast unit 2 Tray sheet metal (base plate) 3 Vapor flow 4 Initial opening limitation (closed position) 5 Lower ballast unit 6 Maximum opening limitation 7 Initial opening area 8 Maximum opening area 9 Centripetal vapor exit

0

Kittel tray with downcomer2)

Expanded-metal tray as previously described with controlled liquid flow 1 Expanded-metal segment 2 Channel 3 Downcomer 4 Inlet weir

3

References

P.21 [2.74]

2.5 Countercurrent Distillation (Rectification)

179

Tab. 2-19. (continued) ~~

Tray Explanatory sketch

Design and operation principle

Jet tray

Cross-flow tray with half stamped out slots. The flap opening acts as a nozzle. The vapor jet forms parallel to the tray and forces the liquid to flow across the tray to the downcomer. To avoid draining of the tray a minimum vapor velocity is required.

[2.74]

Cross-flow tray made of expandedmetal segments. Vapor outlet aperture - vapor blow direction - are shifted by 90". To increase the active surface of the tray and to avoid entrainment of liquid baffle expanded-metal deflectors are mounted in a sharp angle against the vapor flow direction. 1 Baffle 2 Downcomer 3 Upper tray downcomer and weir

v.21 [2.88]

Cross-flow tray for vacuum operation. Flat plates including downcomers are arranged at a small distance. The liquid flows as a film across the plates in constant contact with the vapor. Liquid and vapor both flow through the downcomer to the next tray.

12.21

_ -

t

' \__ >._ /'

Perform contact tray5)

1

2

3

Film tray

') *)

3, 4,

5,

Representation Representation Representation Representation Representation

References

according according according according according

to to to to to

ACV Dr. Stage, Schmidding, Heckmann, Koln-Niehl. Stahl GmbH, Viernheim. Montz GmbH, Hilden. Kiihni, Allschwil. Gutehoffnungshutte AG, Oberhausen-Sterkrade.

180

2 Distillation and Partial Condensation

Table 2-20. Geometrie and fluid dynamic characteristics of important column trays [0.1, 0.6, 2.73, 2.741. Tray type, sizes (mm) and operating variables

Vacuum rectification

Normal pressure operation

Overpressure rectification, absorption

500- 8000 80- 160 1.25 . d, (0.5-0.6) . d 20-30 0.7. h, 500-800 (0.7-1.2) * ego.’

500- 8000 80- 160 (1.25-1.4). d, (0.6-0.75) . d 30-70 0.8 * h, 400- 600 (0.7-1.5) * ego.’

500- 8000 80- 160 1.25 . d, 0.85 ’ d 40-100 0.9. h , 300-400 (0.6-1.4) . ego.’*)

(8.4- 12.5) . w

(10-16)

0.05-0.15 0.20-0.30

0.10-0.30 0.35-0.75

0.15-0.40 0.55- 1

500-8000 80- 150 300- 600 20- 30 500- 600 (0.8- 1.2) . e[O.’

500-8000 80- 150 300- 600 40- 60 400- 600 (0.7-1.5). Q[O.’

500- 8000 80- 150 300- 600 60- 80 300- 500 (0.6-1.4) . ego.’

Bubble-cap trays Tray diameter d Cap diameter d, Cap spacing t (division) Weir length I , Discharge weir height h , Weir outflow height h, Tray spacing Az Gas(vapor) velocity w (m/s) Gas(vapor) velocity in chimney Weff (m/s) Pressure drop A p dry (kN/m2) total (kN/m2)

(5.5-8.4)

*

w

.w

Channel trays Tray diameter d Channel width Channel length Discharge weir height h , Tray spacing Az Gas(vapor) velocity w (m/s) Gas(vapor) velocity in chimney weff (m/s) Pressure drop A p dry (kN/m2) total (kN/m2

(6.3-8.4)

*

w

(8.4-10)

.w

(10-12.5)

.w

0.10-0.1 6 0.22-0.30

0.15 -0.30 0.35-0.60

0.20-0.40 0.45-0.70

500-10000 50- 150 820 2040 1.5 . d, 0.22-0.32

500 - 10000 50- 150 820 3050 (1.7-2.2) . d, 0.16 - 0.24

500- 10000 50- 150 820 4070 (2-3). d, 0.12-0.16

500-800 (8- 14) * ego.’

400- 600 (10-18) ego.’

300-500 (8-14) . Q ~ O . ’

0.20-0.30 0.30-0.40

0.30-0.40 0.40-0.65

0.35-0.50 0.50-0.90

Valve trays Tray diameter d Valve diameter d, Valve lift Discharge weir height h, Valve spacing (division) t Valve arealtray area (scale for active tray area) Tray spacing Az Maximum gas (vapor) velocity in the openings (m/s) Pressure drop Ap dry (kN/m2) total (kN/m2)

2.5 Countercurrent Distillation (Rectification)

181

Table 2-20. (continued) Tray type, sizes (mm) and operating variables

Sieve trays Tray diameter d Openings (drill hole) diameter dd Hole spacing (division) t Opening areahray area (opening ratio) Discharge weir height hw Tray spacing Az Minimum gas (vapor) velocity in the openings w e f f , m (m/s) Gas(vapor) velocity in the openings weff (m/s) Pressure drop dry (kN/m2) total (kN/m2)

Vacuum rectification

Normal pressure operation

Overpressure rectification, absorption

500- 4000 2.5-15 (2.5-3). dd 0.12-0.20 10- 20 500-800

500-4000 2.5- 15 (3-4) . dd 0.08-0.15 20- 50 400- 600

500-4000 2.5-15 (3.5-4.5) . dd 0.06-0.10 40- 80 300-400

10 *

g

10

ego.’

eg0.5

. @p0.5

1.8 . Weff,m

1.8 . Weff,m

1.8 . Weff,m

0.08-0.20 0.20-0.30

0.08-0.30 0.35-0.60

0.08-0.35 0.55-0.90

500- 4000 3- 12 0.25-0.35

500- 4000 3- 12 0.15 -0.25

500- 4000 3- 12 0.12-0.18

Grid trays Tray diameter d Slit width Slit arealtray area (opening ratio) Liquid level on the tray Weir height Tray spacing Az Gas(vapor) velocity w (m/s) Ratio of minimum gas and vapor velocity in the slits w e f f , m (m/s) Pressure drop Ap dry (kN/m2) total (kN/m2) *)

8- 10 12- 28 30- 50 no weir existent; no guided liquid flow 400- 600 300- 400 250- 300 (1.5-2.3). ego.’ (1.1-1.6) ego.’*) (1.4-2) . Q;’.’ 10 . ego3

10 . ego5

10 *

0.06-0.10 0.10-0.20

0.08-0.12 0.20-0.40

0.08-0.12 0.40-0.65

eg Gas (vapor) density (kg/m3)

Fig. 2-61. Explanation sketch for Table 2-20. Section 1-11

Az d

Active tray area Outlet downcomer area Inlet downcomer area Spacing Weir length Discharge weir height Weir outflow height Tray spacing Tray diameter

Section 1-11

--t

Gas (vapor) p h a s e phase

- - C Liquid

eF0.5

182

2 Distillation and Partial Condensation 0

0

-ES -LB

0

-BS

0

Fig. 2-62. Sections of a cross-flow tray. WA Working area IDA Inlet downcomer area LDA Liquid distribution area DA Disengaging area (outgassing area) ODA Outlet downcomer area DS Dead section BS Bubbling or froth section ES Entrainment section LB Liquid backup in downcomer to form a gas(vapor) seal T Gas, vapor 1 Liquid

Minimum vapor load (weeping, draining): if the vapor load is too low, instability of vapor distribution on the tray and weeping, leaking of liquid through the tray occur Minimum liquid load (minimum crest over weir, weir overflow): if the liquid load is too low, a constant weir overflow and a perfect liquid distribution on the tray are not guaranteed Maximum vapor load (entrainment, flooding): if the vapor load is too high, liquid droplets are entrained in a considerable amount or even flooding occurs Maximum liquid load (downcomer blockage, foam disintegration): if the liquid load is too high, the downcomer capacity is not sufficient and the too low residence time in the downcomer does not allow a complete separation of vapor and liquid (foam disintegration)

A detailed overview of the dependency of selected trays on the operating range, the tray geometry and the operating parameters, is given by MOLZAHN and SCHMIDT in [2.92]. For example, an increase in the tray spacing gives an increase in all loading limits. An increase in the free area gives an increase in the vapor load limit, while the liquid load limit remains unchanged. Increasing the height of the weirs expands the operating range of the liquid, but reduces the vapor load range. An increase in the downcomer height shifts the operation range to higher liquid loads, but decreases the range altogether. The maximum possible vapor load F, or wg,, is reached under a given liquid load when the height of the two phase layer on the tray of interest is equal to the distance between the two trays. A liquid droplet of diameter d , is then suspended between the two trays. The weight of the droplet, reduced by the buoyancy, is then equal to the resistance of the flow.

183

2.5 Countercurrent Distillation (Rectification)

Table 2-21. Comparison of some important column trays [2.74]. Tray type

Operating range Crnax/Grnin

5

4 6 8 8

Flexi- Total bility pressure (Vo)*) drop of the tray at 85% of the maximum loading (mm H2O) (10 Pa)

(-1

at 85% of the maximum loading

(->

Bubble cap tray(+) 4 to Channel tray (+) 3 to Thorman tray 4 to Valve tray (Koch) 5 to Valve tray 5 to (Glitsch) (+) Sieve tray (-) 2 to Kittel tray 2 to Drain sieve 2 to tray ( +1 Turbo grid 1.5 to tray ( + 1

Tray efficiency in the allowable range of loading changes

0.8 0.6 to 0.8 0.6 to 0.7 0.55 to 0.65 0.85 0.7 to 0.9 0.8 0.7 to 0.9 0.8 0.7 to 0.9

80

50 80 80 80

45 to 50 to 45 to 45 to 40 to

80 85 60

60

0.7 0.7

0.7 0.6 0.5

300 to 400 300 to 500 300 to 500

0.5

300 to 500

60

40

3

0.7 to 0.8 0.7 to 0.8 0.6 to 0.8

10

30 to 50 20 to 50 30 to 40

2.5

0.7

0.6 to 0.8

10

25 to 40

3

55

(-)

900 to 800 to 400 to 400 to 400 to

0.8 0.8 0.75

3

Equiva- Weight lent cost (N/m2) related to the bubble cap tray

1

0.8 0.8

1400 1400 700 700 700

Tray works well with polluted liquids Tray works badly with polluted liquids Flexibility, the part of the operating range in which the tray efficiency exhibits only a +15% swing G,,, Gas (vapor) loading at the top limit of the operating range Gmin Gas (vapor) loading at the lower limit of the operating range

(+) (-) *)

from Eq. (2-93) ~OIIOWS (2-93) where w , , ~is the maximum vapor velocity referred to the tray area and diameter d, of the droplet. Droplets with a diameter < d, are carried out; droplets with a diameter > d, fall back into the bubble regime. c , ~ is the frictional coefficient (c, = 0.4 with turbulent flow past immersed single drops). Introducing the F-factor

F,,,

=

3 * c,

(2-95)

where F, is the maximum F-factor to the droplet diameter d,. SOUDERS and BROWN [2.93] include the first root expression as the loading factor k,. This depends on tray construction, tray spacing, flow pattern of the phases, properties of the mixture, and particularly on the size of the droplets which is influenced by the surface tension. The loading factor k, has to be found ex-

2 Distillation and Partial Condensation

184

I

I

I

Flooding

rc.; .....................r........

Downcomer

ition

/

-

Weepinq, Draining

[m3/(mZ h)l

Fig. 2-63. Tray columns with guided liquid flow, range of operation [2.14, 2.921. Aps Height or stage specific pressure drop 7

v, Liquid loading wn Vapor velocity

perimentally and is usually given by the 80% vapor load related to the flooding manufacturer of the trays. Figure 2-64 al- point; the column diameter d, the number lows a first approximation of k , for sieve of exchange elements n and the outlet and bubble cap trays. For valve trays, k , may be assumed to be 30% larger. Empirical approximations are given in Table 2-22 to calculate the loading factor k , for important types of trays. The column diameter d can be computed by Eq. (2-92), after the velocity w related to k,Q.OL the free cross-sectional area AQ is calcu- [rnlsl lated with F, in Eq. (2-95) and a safety allowance is added. An alternative method by STICHLMAIR 0 0 2 [2.97] to compute the column diameter d and the tray spacing Az is explained in I I / Table 2-23. Since the loading factor k ,is 200 300 LOO 600 800 1000 m m not used, the use of this method is advanAz Irnml tageous. In practice, design charts are sometimes Fig. 2-64. Load factor k, for sieve and bubble given by the column manufacturer. Hence, cap trays with weir height of 30 mm. Representawith the charts, as shown in Fig. 2-66 start- tion according to MERSMANN [0.1], Vol. 2. ing with the density ratio Q,/,Q,, tray spac- Az Tray spacing ing Az, volumetric vapor flow Vg and an k, Load factor

t

2.5 Countercurrent Distillation (Rectification)

185

Table 2-22. Empirically fixed estimates for the calculation of Souders, Brown correlation factors k , for the upper loading limit of selected column trays [2.2], [2.95]. 0

Souders, Brown correlation factor for the upper loading limit of tray columns

F w

eg, el

0

F Factor (see for example [O.l]) Gas or vapor velocity, related to the free column area (m/s) Gas or vapor density, liquid density (kg/m3)

Souders, Brown correlation factor for the upper loading limit of bubble cap trays k , = 0.05

Az

z,

d,

0

VKZZ,

vz

Tray spacing (mm) Cap height (mm) Cap diameter (mm)

Souders, Brown correlation factor for the upper loading limit of sieve trays

dd 0

.

determined for an opening ratio of 8 % Hole diameter (mm)

Souders, Brown correlation factor for the upper loading limit of Koch valve trays

4

Reduced gas or vapor loading (m3/s) System factor which is characterized by the mixture foam formation tendency (fs= 0.85 for vacuum operation, f,= 0.75-0.80 for normal and overpressure operation) Factor which considers the influence of tray spacing Column diameter (m) D

Vg.r = ~. 3600

ME

-. Qg

1

f , . f ~* , ~

Molar mass of the gas or vapor (kglkmol) Maximum gas or vapor flow rate in the column section of interest (kmol/h)

Reduced liquid loading (m3/h) Reflux or solution flow (kmol/h) (continued next page)

186

2 Distillation and Partial Condensation

Table 2-22. (continued)

Reduced liquid load

?,r

[rn3/hl-

Fig. 2-65. Diagram for the estimation of the diameter d of Koch A type valve tray [2.96] (see also Fig. 2-59).

Table 2-23. Scheme to determine the column diameter d and tray spacing Az by STICHLMAIR 12.971

I

Calculation of the maximum F-factor F,,,: F, = 2.5 . [p2 where a, = a,

AH ~

A ok

01

- (el -eg) g]”4

Relative free cross section

Initial estimation (e.g., a, = 0.1)

AH Hole area, slit area (m2) A,, Active area (m2) O/ Surface tension (kg/sZ)

el e,

g

Liquid density (kg/m3) Gas or vapor density (kg/m3) Gravitational acceleration (m/s2)

(2-96) (2-97)

187

2.5 Countercurrent Distillation (Rectification) Table 2-23. (continued) 0

Calculation of the allowable F-factor and, therefore, fixing of the upper loading limit F=0.7*Fm

0

(2-98)

Calculation of relative gas or vapor velocity wg on the active tray area and calculation of the active area A,,

wg =

F

(2-99) (2-100)

<

0

Gas or vapor volume flow in the column (m3/s)

Calculation of column cross section A and column diameter d (2-101)

where A , Outlet or inlet downcomer area (m2) I, Weir length (m) I,/d Initial estimation (see Table 2-20)

”-0”” 1 - 0.6

(2-102)

(2-103) 0

Calculation of the height of the froth regime height h, which is allowed to be 70% of the tray spacing and determination of the tray spacing Az (2-104)

where h, Height of the two phase zone (froth regime) (m) h , Weir height (m), which is chosen C Constant (C = 0.65 m213) h , ,Weir overflow height (m) 2/3

(2-105)

Liquid volume flow rate (m3/s)

AZ = 1.3 . h,.

(2-106)

Vapor volume flow d

?-

ig1rn3/sl

2

<

n

U

0

c

c

9 2

Fig. 2-66. Nomogram for the determination of the diameter, number of transfer elements and" related shaft area for a cross-flow tray of type KSG-S (see Table 2-19) of Montz GmbH. Representation according to data of Montz GmbH, Hilden.

-

Density ratio f,/jg

E;;

6

I',

Example: given: el/@,= 116, Az = 300 mm, = lo4 m3/h, A , = 0.95 m2, design at 80% vapor loading gives: we = 6.8 m/s, tray inner diameter 2600 with n = 308 transfer elements and about 18 070 downcomer area.

190

2 Distillation and Partial Condensation

downcomer area are determined for a cross flow tray of type KSG-S manufactured by Montz.

Egm =

Yn-Yn-1

Y*(x,)-Y,-l

where

Plate Efficiency

y , - y,,

-,

The mass transfer between gas and liquid phases on a column tray consisting of two phases is usually expressed by the enrichment ratio, plate efficiency E (see Chap- x, ter 1.1, stage efficiency factor). Since the main resistance to the mass transfer occurs in the gas phase during rectification and ab- Y *(Xfl> sorption, E is mainly defined as the efficiency with respect to gas, Eg or Egm. On a column tray, the gas and liquid flow crosscurrently, therefore, one has to distin- y *(x,) - y , guish between the efficiencies Eg and Egm. Eg is a local variable point efficiency on a particular tray with a gas liquid dispersion layer, describing the conditions along a stream line of the gas

Eg =

Y -Yn-1 Y *(XI - Y,,- 1

(2-107)

where y - y, X

Y *(XI

-

the change of concentration of the key component in the gas phase along a stream line concentration of the key component in the liquid phase at that location where the stream line penetrates through the liquid phase equilibrium concentration of the key component in the gas phase related to x

(2-108)

true change of concentration of the key component in the vapor phase from below and above the tray due to the transfer on the tray concentration of the key component in the liquid phase leaving the tray concentration of the key component in the gas phase in phase equilibrium with

-, maximum possible theoreti-

cal change of concentration of the key component in the gas phases leaving the tray when phase equilibrium is reached

Eg and Egmare equal with complete liquid

mixing on the tray. If the liquid is not completely mixed, the concentration x is locally different across the tray. Egmmay be larger than 1 if x > x, and therefore y , > y *(x,). If plug flow of the unmixed liquid is assumed and complete gas mixing below the tray, according to LEWIS22.981, the link between tray efficiency and point efficiency is

L

The enrichment ratio Eg is a true local- From Eq. (2-109) it can be seen that the tray ized efficiency (point efficiency). It can efficiency E,, with no or very poor liquid only be 1 with total liquid mixing in the mixing, is always larger than the point efficiency Eg.A connection between plug flow total depth of the froth regime. Egm is the more important design effi- and ideal mixing in the two phase regime ciency of the whole column tray (MUR- E and Eg for the general case can be obsm tained if the true concentration profile of PHREE efficiency)

2.5 Countercurrent Distillation (Rectification)

191

the liquid x along the pass length is found. If a substance is transferred from the gas From Fig. 2-67 the functional link of Egm phase into the liquid phase and bonded chemically, the mass transfer rate increases. to Eg is then The influence of the reaction on the mass transfer is defined by an enhancement fac(2-110) tor E given in Eq. (3-33). In the case of a chemical reaction superimposed on the mass transfer is The Peclet number Pe is used as a parameter for the liquid mixing, (2-111) where I , is the distance the liquid flows on (2-121) the tray, DE is the dispersion coefficient, and rl is the residence time of the liquid in Overall, the tray efficiency E,,,, is depenthe two phase regime. m is the slope of the dent on the gas and liquid loads of the tray, equilibrium curve, L is the liquid load and the tray geometry, the properties of the G is the gas load of the column. Complete mixed phases and their composition, the mixing of the liquid is noted if Pe = 0, and degree of mixing, especially in the liquid for plug flow, Pe = a. phase, the overflow behavior and the proFigure 2-68 shows a diagram for the cess conditions. Entrainment of liquid by approximation of the point efficiency Eg the upflowing gas phase affects the counteron the gas side, according to STICHLMAIRflow of the phases and leads to axial [2.99]. backmixing of the liquid phase. The effect The point efficiency Eg may be calcu- of the entrainment on the tray efficiency, lated by with the assumption of a small ratio of the gradients of the equilibrium curve and operating line in a simplified form, is accord(2-1 12) Eg = 1 - exp ( - NTU,,) ing to COLBURN [2.101]

if the behavior of the two phase regime and its mass transfer is known. In the case of an absorption and rectification process, the point efficiency is given in Table 2-24. According to STICHLMAIR [2.100], for the bubble regime and approximately for the drop regime of a column tray the following relationship is valid

(2-122)

where Egm is the tray efficiency at an entrainment rate, E. A deterioration of Egm due to the entrainment may be neglected if E < 0.1 - i. Due to the numerous variables mentioned affecting the efficiency it is very difficult to present a valid general approach. For the practical design of a column, experiments under operating conditions in pilotscale plants should be used to find the tray efficiency [2.107].

192

2 Distillation and Partial Condensation

efficiency

STICHLMAIR [2.73].

Normal pressure rectification Vacuum rectification Absorption at a liquid viscosity up to 0.005 Pa s(5 cP)

-

Absorption at a liquid viscosity of 0.01 . . .0.02 Pa. s (10-20 cP) COz-Absorption in water

t

L

I

I

100

1000

02-Desorption from

A' water

0.0001 0.1

10

m.6

;L

L

Fig. 2-68. Diagram to estimate the gas phase point efficiency, by STICHLMAIR [2.99]. Representation according to MERSMANN [0.1, vol. 21. a) N g = 2 ; N , = m b) N6 = 2 ; N / = 15 C) N g z 2 ;N , = 2 d) N g = I ; N ! = l e) Ng = 0.5; N, = 0.5

2.5 Countercurrent Distillation (Rectification)

193

Table 2-24. Determination of the point efficiency E, in absorption and rectification processes [2.73]. Point efficiency Eg E, NTU,,

=

1 - exp ( - NTU,,)

Number of transfer units related to the gas phase 1

NTU,,

m L, G 0

_- 1 -

m

Ng

1

+v.-

L / G N/

(2-113)

Slope of the equilibrium line Liquid and gas flow rates

Gas phase and liquid phase transfer unit Ng, N, Ng = N/ = a

hS A,,

w,

p,, 0

(2-112)

p, . a . h, . A,,

G p/ ' a ' hs

pg . a . hs

-~

w€!

'Auk

(2-114) (2-115)

L

Specific interfacial area Two phase region height Active tray area Gas velocity related to Auk Gas and liquid mass transfer coefficient

Number of transfer units NTU,,, gas phase

(2-116)

(2-117)

(2-118)

Phase contact time (2-119)

Diffusion coefficient, gas phase and liquid phase Height of clear liquid Gas fraction in the two phase or froth region

-...-/I a.

NTU,,

=

4 Dg

hs

wg

Eg

v=

(2-120)

194

2 Distillation and Partial Condensation

Figure 2-69 shows the dependence of the tray efficiency Egm and tray pressure drop Ap, on the vapor load factor F for different types of column trays. With a known tray efficiency E,, as a mean value over the column section of interest, and a previously determined required

number of theoretical stages N,, the number of actual trays Np required in the column is NI Np = -

Ew

(2-123)

10

08 06 €gm

04

02

t

Ap

Imm H,OI

0 8o 60 40

20 0

1

100

80

AP” [mm H,O160

LO

20 O

i

i

i

i

i

i

i

i

i

i

i

i

-

i

i

0 4 0 6 0 8 1 0 1 2 1 L 1.6 1 8 2 0 2 2 2 4 2 . 6 2 8 3 0 f lm/s [kg/m3)”2

Fig. 2-69. Tray efficiency and specific pressure drop of selected trays as a function of the vapor load. Representation according to BILLET[2.2] and Koch Engineering. 1 Sieve tray (open area 12.3%) 2 Sieve valve tray (Koch, type T 9 F, open area 12.5%, weir height 38 mm) 3 Bubble cap tray 4 Kittel tray Column diameter: 800 mm, tray spacing: 500 mm System: Ethyl benzenehtyrene Reflux ratio: 03, operating pressure: 133 mbar A p Pressure drop of a single tray Apn Pressure drop of a theoretical stage F Vapor load ERmTray efficiency

195

2.5 Countercurrent Distillation (Rectification)

Apt is the “dry” pressure drop of the tray, which is defined by

Column Tray Pressure Drop The following discussion of the pressure drop in counterflow columns which is applied to rectification columns is also valid for absorption and desorption columns. An important process parameter is the pressure drop Apg of a gas flowing through a counterflow column. Therefore, under steady-state conditions the pressure P h at the column head and the pressure at the column bottom pb are fixed

where Ap is the mean pressure drop of the individual column trays and Np is the number of trays installed. (The pressure drop of the gas in the free vapor space between the trays is negligible compared to Ap). With a large pressure drop Apg in a rectification column, heat is supplied at the bottom of the column at a comparatively high temperature, while heat is removed at low temperature from the condenser at the top. This may easily be seen from the vapor pressure curve of the column vapor. This is unfavorable from the energy point of view and increases the operating costs. A high pressure drop over the trays leads eventually to a large liquid holdup in the outlet downcomer, and disturbs the desired performance of the tray. On the other hand, a high tray pressure drop occurs with large turbulence in the two phase region, where mass and heat transfer take place; in this case mass and heat transfer are improved. In general, the choice of the type of tray and the liquid and vapor loads results from a compromise between good mass and heat transfer and an acceptable pressure drop. The pressure drop of a tray Ap usually consists of three parts AP

= Apt -k

AP*

f

APst

(2-125)

(2-126) where c, is the frictional coefficient dependent on the tray geometry, eg is the vapor density and wesf is the actual vapor velocity in the openings of the tray. Ap, is the pressure drop which results from the formation of bubbles by the gas. This essentially depends on the surface tension 0,of the liquid and is (for slow bubble growth on a sieve tray, with a sieve hole diameter db) (2-127) Generally A p , is small compared with A p and is usually negligible. Apst is the hydrostatic pressure drop. This is due to the gas flowing through the two phase regime, or froth section, of height h, on the tray APst

= erg

a

g * hs = hr

. el

*

g = h, * E l

el

*

g

(2-128)

where density of the two phase regime and liquid phase height of clear liquid and two phase regime relative liquid fraction in the two phase regime An additional pressure drop Ap of the tray occurs when a considerable amount of the liquid E is entrained by the vapor due to a high vapor load. The pressure drop is not only caused by the hydrostatic pressure of the liquid, but also by the acceleration of the entrained liquid. Therefore, the term Apb is added [2.73] to Eq. (2-125):

196

2 Distillation and Partial Condensation

Table 2-25. References for the precalculation of the tray pressure drop for important tray types.

Table 2-26. Comparison of different packed tower characteristics.

Tray type

References

Sieve tray Bubble cap tray Channel tray Valve tray

[2.73, 2.102, 2.1031 [2.73, 2.74, 2.82, 2.1041 t2.73, 2.74, 2.861 t2.73, 2.74, 2.82, 2.87, 2.105, 2.1061

Packed tower

Ratio d/d, or dJd,

Ratio Z/d o r Z/d,

Packed column

d/dp > 10

Z/d = 5

Packed tube

4 < d,/dp < 8

20 < Z / d , < 30

Parallel column

d,/d,>

Z/d, = 20

10

d Inner column diameter

where G is the vapor flow, Aw, is the change of velocity of the liquid during the entrainment by the gas, from w I= 0 to W I = wg. Due to the numerous variables, such as tray geometry, substance properties of the mixture phases in contact, gas load, liquid load, and operating pressure, a general equation to calculate the tray pressure drop Ap has not yet been developed. In most cases, the tray pressure drop for the design of a column must be found experimentally. Some literature is listed in Table 2-25 to calculate the pressure drop Ap for important types of trays. In Table 2-21 (Chapter 2.5.6.1) estimated data for Ap for different types of trays are listed. A tray column is flooded if the liquid level in the outlet downcomer is approximately equal to the distance between two trays. Then the tray pressure drop is the hydrostatic pressure of the liquid level in the outlet downcomer. The maximum liquid load which leads to flooding of a column may be derived with this [0.1, vol 21. 2.5.6.2 Random Packing, Packing with Regular Geometry

To increase the phase surface area in packed towers (Fig. 2-70 and Table 2-26) randomly packed filling material, and regular or stacked packing made of wire, expanded

dp Typical filling elements dimension d, Diameter of packing (= d of a packed column) Z Height of packing elements

metal, sheet metal, or braided metal are used. Gas and liquid phases still remain coherent during contact. Due to gravity, the liquid flows down and forms a wetted liquid film on the free packing surface. Gas flows over the liquid film. Therefore, the surface area, important for mass and heat transfer, corresponds to the surface area of the liquid film. Fig. 2-70. Packed towers. b Representation according to Raschig GmbH, Ludwigshafen and Sulzer AG, Winterthur. a) Random packed tower b) Regular or stacked packed tower LD Liquid distributor BL Bed-limiter PS Packing support LR Liquid re-distributor SU Support LC Liquid collector RC Ring channel VI Vapor inlet CB Column bottom CL Circulation line to evaporator MH Manhole FO Foundation LS Column support

2.5 Countercurrent Distillation (Rectification) 0)

Vapor out

b)

197

VoDor out

MH

MH

MH

in

CL LS

FO Bottom product

Liquid out

198

2 Distillation and Partial Condensation

Packed towers are used for rectification under vacuum conditions, for gas scrubbing and absorption, for vaporization and in liquid-liquid extraction. Applications are also direct heat transfer between gas and liquid phases and/or between gas and liquid in heterogeneous reaction processes (often with parallel phase flow). In comparison to tray columns, packed towers have a low specific pressure drop per unit height or equivalent stage, a lower liquid holdup, a lower residence time and a higher loadability. Also, geometrically simple packing elements can be manufactured from most materials (steel, stainless steel, copper, carbon, earthenware, china, glass, plastics, etc.) and therefore can provide the possibility of processing corrosive substances. To obtain the most effective exposed wettable surface in a column with random packing (Fig. 2-71), the liquid should be

GO

GO

BL

BL

1

1

z

i.1

LO

z

t

*O

SP

SP

u

LO

Fig. 2-71. Packed towers. PC Packed column TT Tower with packed tubes G1 Gas inlet GO Gas outlet LI Liquid inlet, liquid distributor LO Liquid outlet A1 Heating or cooling agent inlet A 0 Heating or cooling agent outlet SP Support plate BL Bed-limiter

t

evenly distributed and hence the whole packing surface wetted. Since the relative void fraction E (porosity) of the packing near the wall is larger than in the center, the downflowing liquid tends to preferentially trickle down the wall with increasing distance from the liquid distributor. To restrict this specific effect of maldistribution, liquid redistributors are used to interrupt the packing. Another reason for redistributing the liquid several times, is because of so-called channeling in packed columns. Liquid flows in individual channels through the packing and the packing surface is only partially used. With vacuum operation and small liquid loads, channeling is particularly observed. During start up of a column, total wetting is obtained by complete flooding of the packing before the desired operation loading is reached. To restrict maldistribution and channeling of the liquid, the ratio of the packing diameter d or d,, and the characteristic diameter of filling material dp, respectively should be larger than 10. For separation processes with heat supply or withdrawal, towers with packed tubes are employed (Fig. 2-71). These feature tube bundle apparatus with random packing in the tubes and a cooling or heating medium pumped through the shell. For difficult separations in small-scale plants, ‘parallel columns”, towers with packed tubes, including miniature high performance packing, are used. Packing Design, Demand to Packing

To obtain good phase distribution and maximum exposed liquid film surface for mass and heat transfer, different design forms of packing are inserted, for example, spheres, rings with smooth and profiled surfaces, partitioned rings, wire gauze rings, saddles, and spirals. Table 2-27 gives

199

2.5 Countercurrent Distillation (Rectification)

Table 2-27. Common random packing element shapes, characteristics, and design data. Element type

Material

Raschig rings')

ceramic

carbon steel V2A,V4A

stainless steel

Pall rings')

ceramic metal plastic

w

Dimension (length outside x diameter x wall thickness) (mm)

Bulk No. of density elements (kg/m3) per m3 (1000/m3)

Specific Percent surface void area space a (m2/m3) E (0101

210 46 6.4 0.75 230 51 6.5 0.75

330 195 98 44 350 220 110 48

70 73 78 81 92 93 95 96

620 550 510 460 5 0 ~ 0 . 8 320 120 1 5 x 1.0 85 25x 1.3 72 sox sox 1.5

46 6.3 220 52 6.3 215 52 6.4

220 120 3 60 215 105 350 220 110

73 78 93 94 96 88 90 92

15x 15x2

25 x 25 x 3 50x 100 x 15 x 25 x 50 x 100 x

50x5 100 x 10 15 x 0.5 25 xO.8 50 x 1.0 100 x 1.5

25 x 50x 15 x 25 x 50x 15x 25x

25 x 3 50x5 15 x 0 . 4 25 xO.6

700 620 520 450 660 640 430 300

Intalox saddle')

ceramic

15 25 50

670 610 530

400 85 9.3

450 255 120

71 74 79

Berl saddle')

ceramic

15 25 50

800 700 600

280 75 8.0

430 260 120

67 69 73

(continued next page)

200

2 Distillation and Partial Condensation

Table 2-27. (continued)

Element type

Material

Dimension (length outside x diameter x wall thickness) (mm)

Super-Torus saddle2)

plastic

hominal size

Super saddle')

1 2 3

plastic 25 x 25 (PolYProPYlene) metal 25 x 25 xO.25

Bulk No. of density elements (kg/m3) per m3 (1000/m3) 80 57 40

40

118

Specific Percent surface void area space a (m2/m3) E (VO) 240 110 90

90 94 95.5

90

258

95

190

95

262

95

520 211 56.5

360 260 160

93 95 94

6.2 1.2

Interpack filling')

metal

1 5 x 15 x0.4 2 0 x 20x0.4 3 0 x 30x0.6

460 350 330

Top Pack')

metal

80 x 0.6

170

75

98

Hacketten')

plastic

45 90

63 53

135 108

93 94

201

2.5 Countercurrent Distillation (Rectification)

Table 2-27. (continued) Element type

Material

Dimension (mm)

Bulk No. of density elements (kg/m3) per m3 (1000/m3)

Specific Percent surface void area space a (m2/m3) E (VO)

Hedgehog')

plastic

40 56

124 106

300 184

86 88

VSP')

metal

25 40

205 110

97.5 98

Hiflow Rings3)

metal

35 50 25 90 20 75

130 93 218 59 285 61

98 98 92 96 76 85

plastic ceramic

')

2, 3,

170 170 75 35 560 345

11.2 4.95 45.5 1.25 118.7 2.08

Representation according to Vereinigte Fullkorper-Fabriken GmbH + Co, Ransbach-Baumbach, Representation according to Raschig GmbH, LudwigshafedRhein. Representation according to Rauschert GmbH + Co. KG, Steinwiesen.

202

2 Distillation and Partial Condensation

an overview of design, characteristic data and dimensions for several important packing types. The following characteristics for the selection of packing should be considered: High separation effect High loadability Low pressure drop Ability to compensate existing phase maldistributions and limited side phase mixing Little tendencies to cause channeling and poor wall effects on flow (maldistribution); good wettability Sufficient mechanical strength with regard to pressure and impact Easy cleaning Low manufacturing cost The demand of high separation with a simultaneous low pressure drop is particularly important. The separation efficiency and pressure drop of packing increase and loadabilty decreases with decreasing dimensions of the packing material. The dependence of the separation effect on the loading is more pronounced for packing material with small dimensions than with larger dimensions. With constant conditions, the separation effect may be increased by a higher liquid load.

Packing with Regular Geometry The geometry of regular packing [2.144] with a defined controlled flow manner, offer an even phase distribution across the column cross section and increased phase turbulence with careful liquid distribution. Generally, maldistibution does not occur. Since the separation effect is independent of diameter and height of the packing, scale up from experimental scale to technical scale may be carried out easily. Table 2-28 gives an overview of the regular packing used in process practice.

Regular structured packing offers the following advantages in comparison to other packing:

0 0

Higher loadability and better separation effect Lower specific pressure drop Less required packing volume and therefore a lower height required for mass and heat transfer

Regular packing is mainly used with vacuum rectification and absorption, especially when the column pressure drop is limited.

Loading Limit, Operating Range The operating range is given by the loading limits which are dependent on the type and geometry of the packing, and on the properties of the counterflow phases. Figure 2-72a shows a typical operating range of a packed column (Fig. 2-72 b). The upper load limit (flooding boundary) must not be crossed under any circumstances. If this limit is reached with a high vapor and liquid load, the countercurrent flow of liquid and vapor breaks down. The liquid is build up in the packing by the vapor and a froth layer of vapor and entrained liquid can arise. Therefore, beyond the upper limit a fast decrease of the separation effect and a steep increase on the pressure drop occur. (Under flooding conditions, the packed column acts as a bubble column with packing. A further increase of the gas velocity fluidizes the packing material, the operating range is now like a fluidized washer.) The minimum liquid load (diswetting limit) is the liquid load when the packing is just evenly wetted. Therefore, below the lower limit, the exposed wetted surface and with it the separation effect quickly decrease. The best separation effect is obtained close to the upper load limit, in the “emul-

8. Some selected arranged type packings with regular structure.

packing Characterization Design Material

with ucture

Favored application

Porosity (-)

0.750.98

Falling film flow, good surface wetting

Specific Gas load surface factor (m2/m3) (1/pa)*)

Liquid No. of theoretical flow rate separation (m3/ stages per m **) (m2 h)) *)

Sp pr (m

60-700

Rectification under vacuum up to ca. 0.5 mbar

400

1-5

to 10

2.5-5

0.

trickle Montz ) A3-500

Braided metal packing Rectification where single packings are under vacuum shifted by 90" against up to each other ca. 1 mbar

500

0.6-2.5

to 15

4-7

0.0

B1

Packing made of 0.2 mm embossed metal sheet with special surface structure and round form flow channels

100-500

0.6-3.0 Bl-300

to 30 B1-300

ca. 4 B1-300

0. B1

C1

Similar to type B1, but made from PTFE or polypropylene

raided Spirally wound thin wire packing (diameter 0.2-0.3 mm) gauze, arranged to ss packings up to ca. 0.3 m z with spacers in between )

g r AG)

Rectification under overpressure, normal pressure and vacuum Absorption

Packing material is made from diagonally folded lamellae and arranged to provide open crossing channels, diagonal to the column axis. Successive individual shifted packing heights are 160-300 mm

(continued

8. (continued)

packing Characterization Design Material

Favored application

Porosity (-)

Specific Gas load surface facsr (m2/m3) (V'Pa ) * )

0.96

Rectification Absorption

Packing material is made from reverse corrugated metal sheet perforated with slots and arranged to form open crossing channels at 45" to the column axis. Packing height is 250 mm.

0.75

Rectification of corrosive mixtures in glass columns or enameled columns Absorption

Thin-walled Al-silica based ceramic

k

250

0.97 Rectification in the range from vacuum to overpressure Absorption

Corrugated sheet metal packing of up to 11 m column diameter

ak

500

Rectification in vacuum starting at 1 mbar up to atmospheric pressure

Metal gauze packing

X

er g ak ig )

0.9

450

250

0.6-3.2

2.0-3.0

1.0-2.5

1.0-3.0

Liquid flow rate (m3/

(m2 . h)) *) >0.2

0.6-200

>0.2

0.6-200

No. of theoretical separation stages per m **)

4-6 (ca. 5 at 75% of the Splashlimit 2.4 V'Pa) 2.5-3.0

4-5

3-3.5

er g pak i AG)

Packing is made from grids arranged in layers with diagonal thin metal sheet lamellae with smooth or rippeled surface

Rectification Absorption

Rectification acking Packing consists of sets of two hollow hert Absorption triangular pyramids ) connected at the vertex (forming a type of “hourglass”) with further constriction inbuilt (see Table 2-18) Continous surface renewing, less good surface wetting

with cture

Ceramic packing

yp 100

g acking grid g GHH N AG

Horizontal and crosswise layer made from folded and punched out sheet metal or expanded metal

0.81 0.87

0.87

230

1.0- 3.5

2.5-3.5

60- 110

60

1.0-3.5

up to 150 ca. 1.5

Rectification Absorption

erform- Packing is made from cking expanded metal with shaped perforated holes and arranged horizontally and vertically to provide changing direction of the holes

1.0-3.5

3-ca. 80

0.8-1.5

oximation. oximation for common loading range, operating pressure, and test mixtures.

206

2 Distillation and Partial Condensation

a1

F-Factor

-

Fig. 2-72. Packed column operation range (a) (representation according to MOLZAHNand SCHMIDT [2.92]) and schematic representation of the separation effiency as a function of the vapor load (b). System: Methanol/Methanol-vapor, p = 1 bar, L9 = 64"C, metal Pallrings, packing diameter d,= 1.5 m Gas load factor at flooding limit Gas load factor at allowable upper load limit No. of theoretical separation stages per m Height of packing equivalent to a theoretical separation stage Operating range Gas velocity Liquid flux

2.5 Countercurrent Distillation (Rectification)

207

sification range”, between the flooding empirically determined flooding curve can point and the phase return point. In the op- be calculated from the following: erating range, overflow of gas hinders the liquid; the gas starts to hold the liquid up 0 The direct dependency of the flooding factor, F F / v c on the flow paraand is dispersed in it. meter L / G I/e,/el- (Fig. 2-73) The maximum allowable gas velocity w ~has, to ~be known ~ ~ to determine the di- 0 The dependency of the related pressure drop of the dry packing on a related ameter of the packing d, and, therefore, the sprinkle density (see Fig. 2-76) inner diameter d of the column by means of Eq. (2-92). It has to be chosen in such a way that the column will not be flooded and the (where FF is gas load factor at flooding best possible separation effect is also limit, L and G are the liquid and gas flow rates, respectively, and el,.eg are the liquid achieved. Calculations for the flood loading of and gas densities, respectively.) packed columns are given by SHERWOOD According to BECK[2.118] the maximum , may ~ ~be itera~ [2.112], ECKERT[2.113], BILLET12.21, MERS- allowable gas load w ~= w0 MANN [2.114], SCHMIDT [2.115], REICHELT tively estimated by Eq. (2-131) in Table 2-29, [2.110], ZECH12.1161 and VOCT12.1171. The where a liquid flow i is first assumed.

-

5 2

1

1

2

3

L

-r

5

6

8

is,-

G

3 L

10

15

20

30

LO

SI

Fig. 2-73. Flooding limit for different packings. Source: BILLET,MACKOWIAK [2.119], [2.120]. System: Aidwater; p = 1 bar, T = 293 K 1 Pall rings 50 mm, plastic (polypropylene PP) 2 Hiflow ring 50 mm, PP, 6280 no. pieces per m3, bed diameter 0.3 m, bed height 1.4 m 3 Hiflow ring 25 mm, PP, 45500 no. pieces per m3, bed diameter 0.3 m, bed height 0.9 m 4 lntalox saddle 50 mm, PP 5 Hiflow ring 90 mm, PP, 1340 no. pieces per m3, bed diameter 0.45 m, bed height 2 m 6 Ralu-Pak 250 YC, metal sheet, packing diameter 0.45 m, packing height 1.8 m

208

2 Distillation and Partial Condensation

Table 2-29. Calculation of no. of theoretical stages per m packing height n,, max. allowable gas ~ column ~ head of the random packing and specific pressure drop Ap, by velocity w ~= w,, at~ the BECK [2.118]. 0

Number of theoretical stages per m packing height n, (2-130)

w,

Allowable gas velocity at column head, m/s

K, , K2Specific filling material constants

el, eg

Liquid and vapor or gas densitiy, kg/m3 Bed height, rn 1 Liquid load or flux, m3/(m2 . h), lower limit d Diameter of the packing, m dp Characteristic filling material dimension, m f(d,/d) = 1.72/[1 + 7.5 * (d,/d)2] - 0.72 2,

0

Maximum allowable gas velocity wo at column head, m/s (2-131)

- 'fi

(fr (11~) = 0.775 + 0.225 viscosity function kg/(m . s)) (to be considered at qI > K , , K4 Specific filling material constants qI Liquid dynamic viscosity (cP P kg/(m. s)) 0

Specific pressure drop Apf per m packing height, given in bar (2-132)

f(i. o/) =

+

K7

'

i) .A(%)

K , , K 6 , K7 Specific filling material constants w Actual gas phase velocity, m/s n = n ( i ) (see Fig. 2-74) 0

Specific filling material consants K , - K7, and lower and upper liquid load ConType and dimension Stants, Liquid Raschig Raschig Raschig load ring ring ring ceramic ceramic metal 25 x 50x 25 x 25 mm 50 mm 25 mm KI

KZ

K3

K4

K5 K6 (7

I," ImaX

0.30 0.10

0.062

0.30 0.08

0.083

0.095 0.077 13.5 27 0.115 0.090 0.0145 0.008 8.0 10.0 100 150

Raschig Pall Pall ring ring ring metal ceramic ceramic 50x 25 x 50 x 5 0 m m 25 mm 5 0 m m

Pall ring metal 25 x 25mm

Pall ring metal 50x

50 mm

Berl saddle ceramic 25 mm

Berl saddle ceramic 50mm

lntalox Intalox saddle saddle ceramic ceramic 25mm 50 rnrn

0.32 0.40 0.60 0.45 0.30 0.25 0.50 0.50 0.40 0.30 0.10 0.10 0.08 0.10 0.08 0.10 0.08 0.08 0.10 0.08 0.088 0.078 0.100 0.113 0.075 0.093 0.075 0.093 0.070 0.100 0.077 0.093 0.078 0.090 0.072 0.071 0.058 0.065 0.078 0.065 27 13.5 16.0 9.0 9.0 3.8 13.5 13.5 6.3 6.3 0.115 0.090 0.090 0.090 0.070 0.070 0.110 0.080 0.110 0.080 0.013 0.007 0.0090 0.0068 0.0079 0.0062 0.015 0.0070 0.0115 0.0070 8.0 10.5 13.0 10.5 13.0 10.50 13.0 10.5 13.0 10.0 105 I155 115 180 185 280 145 220 145 220

iminlower liquid load limit (m3/(rn2. h)). ima, upper liquid load limit (m3/(m2 . h))

at an allowable gas velocity of 0.1 m/s.

2.5 Countercurrent Distillation (Rectification) Table 2-29. (continued)

0

Investigated range of Eqs. (2-130) - (2-132) Value

Investigated range

d

< d < 1.20 m 1.0 < Z < 6 . 0 m 400 < < 1400 kg/m3 0.015 < eg < 25.0 kg/m3 0.10 < ‘I[< 90 CP 0.83 < I < 320 m3/(m2 h) 0.15

z

ec Qg

‘I/

i

+

Determination of the exponent n as a function of the liquid load

Z.

300

I-

250

200 1 [m3/(rn2. h)] 150

100

50.

0.

I

0.7

0.8

0.9 n-

1.0

1.2

Fig. 2-74. The curve is valid for metal Raschig rings and metal Pall rings, plastic Pall rings and ceramic saddles I Liquid load n Exponent

209

210

2 Distillation and Partial Condensation

Packing Pressure Drop

Figure 2-75 shows qualitively the dependency of the height specific pressure drop, Aps = A p / Z of the gas phase and the related liquid contents . &) on the gas velocity wE and the cross-sectional area related sprinkle density, B = v//AQ,with counterflowing phases (where and V,are the volumetric flow of the liquid phase and the packing volume, E is the relative void fraction and A Q and Z are the packing cross-sectional area and height, respectively). The following flow areas are observed:

v//(V,

c

q

I

v/

0

0

0

0

Area Z: no mutual influence of both phases, the liquid contents are independent of the gas velocity, the pressure drop of the gas is slightly higher than that for a dry packing due to the reduced penetration volume Area ZZ: beginning of mutual influence of both phases Area ZZT: at beginning of stowing, packing loaded by liquid phase (loading point, stow point, lower loading limit), increase of the liquid holdup with increasing gas velocity Area I V : at the upper loading limit of the packing cflooding point, flood limit), liquids stow, change of phase distribution, with a further increase of the gas velocity blowing out of liquid

C

\

*

limit Flooding

m,

m,

Calculation methods for the pressure drop for two phase counterflow in packing are given by ECKERT [2.113], SCHMIDT [2.115], MERSMANN [2.114], REICHELT [2.110] and ~ U T S C H[2.121]. From Fig. 2-76 the pressure drop A p of the wet packing may be determined [0.26] together with the pressure drop of the dry packing Apt 1-&

Qg’w;

Apt=c,.-

-a

2

c3

a-

z

d,,

(2-133)

log wg

-

B Loading limit

Fig. 2-75. Gas pressure drop (a) and liquid holdup (b) for counterflow of liquid and gas in random packings. APS Specific pressure drop

v,. &

Liquid holdup

wg

Gas velocity

2.5 Countercurrent Distillation (Rectification)

21 1

and the dimensionless trickle intensity B *

where g

c,

B"

-' Reg =

-

Fig. 2-76. Pressure drop and flooding point in random packings. Representation according to Rauschert GmbH, Steinwiesen.

e/.g.z B*

accelaration due to gravity (9.80665 m/s2) resistance coefficient (friction factor) of the packing, (obtained from Fig. 2-77) and the Reynolds number Reg of the gas phase

Dimensionless gas load

dp

wg * dP (1 - E) vg

-

(2-135)

particle diameter of packing material

6 * Vp dp = AP

(2-136)

Dimensionless trickle intensity

50

I

cw

10

1

Re,

-

Fig. 2-77. Friction factor of dry random packings as a function of the Reynolds number of the gas flow. Representation according to Rauschert GmbH, Steinwiesen. 1 Spheres 2 Ceramic Pall rings and saddles 3 Ceramic Raschig rings 4 Metal Raschig rings 5 Bialecki rings c, Friction factor of the dry random packing Reg Reynolds number of the gas flow

212

2 Distillation and Partial Condensation

b , A p volume of packing and packing surface area, respectively v,, vg kinematic viscosity of liquid and gas According to BECK[2.118] the specific pressure drop of a unit height of packing may be estimated by the simple approximation (Eq. 2-132) given in Table 2-29. The flow of gas and liquid phases in packings with regular geometry is similar to the flow of trickled liquid and gas inside vertical tubes. For a two-phase flow in Sulzer braided packing, Zocc [2.122, 2.1231 describes in his work a dimensionless flow number, which is similar to the flow number for counterflow of gas and liquid in tubes given by FEIND[2.124]. The flooding point and pressure drop are expressed as a function of these flow numbers, based on

empirical data. In Table 2-30 some relationships are listed to determine the flow number, flooding point and pressure drop for regular structured packing. Figure 2-78 shows the pressure drop and separation effect of selected packing, as a function of the gas load. Separation Effect, Height of Packing

The separation effects of packing and other packing elements are usually expressed by the concept of theoretical separation stages (see Chapter 1.1). HETS or HETP are used for the packing height, which corresponds to the effect of a theoretical separation stage or tray, or their reciprocal value as the variable n, which is the number of theoretical separation stages per m packing height.

bl

a1

N, Z

Il/ml

t

80 60

* PS

7 1 0 [mrn water1

20 0

Fig. 2-78. Separation efficiency (a) and pressure drop (b) of different regular packings as a function of the gas load. 1 Ralu-Pak 250 YC, sheet metal (chlorobenzene - ethyl benzene, p = 133 mbar, d = 220 mm, Z = 1.4 m [2.120]) 2 Montz braided metal packing Type A3 -500 (chlorobenzene-styrene, p = 66 mbar, d = 220 mm, Z = 1.4 m, presented by Montz GmbH) 3 Sulzer metal gauze packing BX (ethyl benzene - styrene, p = 133 mbar, d = 500 mm, Z = 2m [2.127]) n, = N , / Z Number of theoretical separation stages per m packing height Ap,/Z Specific pressure drop per m packing height F Gas load factor

2.5 Countercurrent Distillation (Rectification)

213

Table 2-30. Determination of the flow number R,, flooding point and pressure drop Ap for regular packings [2.122], [2.123].

Flow number R, to characterize the two phase countercurrent flow in packed columns (2-137) 0

Gas phase Reynolds number Re, (2-138)

0

Liquid phase Reynolds number Re,

Re, =

.

.

r;;

(2-139)

A Q .a . sin H . vl

5. 6

Flow rate of gas and liquid phase Cross section and volume specific surface of the packing Ap a Inclination angle of the trickle plane H e,, el, q,, q,, v,, v, Density, dynamic and kinematic viscosity of gas and liquid phase 0

Film thickness 8,of the laminar falling film (Re, < 400) (2-140)

0

Hydraulic diameter dh 4

dh = -

a

- 2 . S,

(2-141)

Total thickness of packing gauze trickled on both sides

a, = a, + 2 ' S, 0

(2-142)

Example: Sulzer metal gauze packing, Type BX and CY. Type

Specific surface a (m2/m3)

BX CY

500 700

Gauze thickness

6, (m)

0.00045 0.00045

Film length 2, (m)') 0.0237 0.0103

Inclination of trickle plane H

("1

67.1 57.4

Maximum flow number Rs,maxZ)

Operation flow number R, -

50 44

37.5 33.0

~~

Presssure drop A p (2-143) (continued next page)

214

2 Distillation and Partial Condensation

Table 2.30. (continued) 0

Friction factor c, Sulzer BX packing3) at normal operating point c,, = 0.331 Sulzer CY packing3) at normal operating point c,

= 0.68

19.3

+ a.d

(2-144)

~

112

+ __ a.d

(2-145)

Z Packing height d Column diameter ') ')

3,

Film length between two packing inversion points where no mixing takes place Flow number at the upper load limit (exceeding leads to packing flooding) Mean deviation *20%.

1

r= The equivalent packing heights, HETS, HETP or n, depend on: 0

0

0

The type, dimensions and surface conditions of the packing elements The properties of the counterflow phases, such as gas and liquid densities, liquid viscosity and surface tension The operating parameters, such as pressure, vapor load and liquid load

With random packing, the degree of maldistribution u [2.2] of the liquid across the column cross section is influential due to wall effects on flow, channeling, etc., u varies in the limits

05u5l

q is then estimated to [2.125]:

u

v+l

m

4 15

(2-149)

where v is the reflux ratio and m the degree of mixing on each theoretical separation stage

):(

2

m = 27.

(2-150)

If the variable (q), is known for a packed column with evenly distributed liquid, the separation effect (n,),,o of the same column with maldistriubtion may be determined by means of Eqs. (2-148)-(2-150): ~

(2-147)

and the degree of separation effect q is defined as the ratio of (n,), , with maldistribution and (n,), = with even liquid distribution

q=- (nt), > 0 (n,),= 0

v

(2-148)

The smaller the reflux ratio, the more the separation effect decreases with maldistribution. To restrict maldistribution of the liquid by channeling or flow at the wall, the total height of the packing Z is separated into single packing heights 2, each with its own support and cover grids, and liquid distributor [2.118].

2.5 Countercurrent Distillation (Rectification)

3 = 2.5 - 3.0 d ' k

-

d

zk

= 5.0 - 8.0 saddles

-3

d

Rasching rings

5.0 - 10.0 pall rings

(2-152)

(2-153)

(2-154)

Calculation of the separation effect in advance is difficult due to the large number of variables. Also, scale up of HETP or n, found by experiments at pilot-plant scale is not easily done due to an increase in unfavorable liquid distribution in a full-scale plant. A method to calculate the separation effect of packed columns is introduced by BILLETand MACKOWIAK [2.126], where the variable n, is given by

AP (2-155)

as a function of the pressure drop Ap of the packing, the number of transfer units related to the gas phase NTU,,, and the stripping or strip off factor 1

L h=m.G

(2-156)

(rn is the slope of the equilibrium curve and L and G are the liquid and gas flows, respectively). A simple estimation is possible by Eq. (2-130), according to BECK[2.118]. For the stripping and enrichment section the variable n, has to be calculated separately, followed by a control calculation for the single packing of height 2,. Figure 2-79 shows the dependency of the efficiency of selected packing elements on

215

the gas loading. The efficiency is greatest in the holdup region (area 111, Fig. 2-75) between the lower loading limit point and the flooding point. The efficiency increases with smaller dimensions of the packing elements and with a larger reflux or larger liquid load, respectively. The required packing height Z for mass and heat transfer between counterflow phases may be calculated in two different ways as discussed in Chapter 1.9.2. After determination of the desired number of theoretical separation stages Nt by numerical or graphical means, and the separation efficiency experimentally under conditions close to the operating conditions and suitable for scale up, the packing height is

Z = N t .HETS = 4 nt

(2-157)

The packing height may also be determined by the NTU-HTU method (see Chapter 1.9.4.1) based on the kinetic theory of counterflow mixture separation. In rectification processes, mass transfer resistance in both phases must be considered. The mass transfer coefficients in both the vapor and liquid phases should be known in order to evaluate HTU The NTU-HTU method is of no importance for rectification in contrast to absorption. This is mainly because no secured data and no valid calculations for the partial mass transfer coefficients are given for wide ranges. Liquid Holdup of Trickled Packing

Liquid holdup is comprised of the static and the dynamic contents. The static contents remain in the stagnant volumes, gussets, etc., of the packing due to surface tension. The dynamic contents are constantly renewed by the liquid flow. If liquid no longer flows on top of the packing, a cer-

216

2 Distillation and Partial Condensation

tain amount of liquid volume remains in the packing due to adhesion. Calculations of the static and dynamic liquid contents are given in the work of GELBE[2.128] and KURTZ [2.129].

-r-----

0

1-

I

I

I

1 F[Pa”’I

I

1

2

15 10

1 7

3 [ rnb a r /m ] 2 APS

1

0.5 0.5

0.7

1

1.5 F[P~”’]

2

3

2.6 Choice, Optimization and Control of Rectification Units For the initial selection of column internals for rectification columns, knowledge of the vapor load and operating pressure is sufficient. With small vapor loads and, therefore, small column diameters, random packings are used close to the lower limit of the coarse vacuum range. If good separation efficiency at low pressure drops - even with large column diameters and a wide operating range - is required, packings with regular geometry are preferred. If a large column throughput requires a large column diameter and the pressure drop with an ambient operating pressure is of secondary importance, trays are used as column internals. For the final selection of column internals, mixture behavior, operating conditions, characteristic performance figures, costs, pressure drop, and other criteria are important. Substantial criteria are listed in Table 2-31. Different evaluation criteria prove to have an opposite influence. For example, the requirement of good separation efficiency leads to a higher pressure drop and higher costs. The selection of column inter-

Fig. 2-79. Separation efficiency and specific pressure drop of selected packings as a function of the vapor load. Representation according to Vereinigte Fullkorper-Fabriken GmbH, Ransbach-Baumbach. 1 Pall ring I” 2 VSP 25 3 Pall ring 1.5” 4 VSP 40 5 Pall ring 2” 6 Top-Pak size 1 (height 45 mm) n, Number of theoretical separation stages per m packing height Ap, Specific pressure drop per m packing height F Vapor load factor System: Isooctane - toluene; operating pressure:1.013 bar; column diameter: 400 mm; reflux ratio: 03

2.6 Choice, Optimization and Control of Rectification Units

Table 2-31. Column internals selection criteria.

0

Product specification (top product, bottom product, sidestreams) Mixture properties Thermal sensitivity, contamination tendency, foaming properties, material properties Operation condition Operation form, operating pressure, reflux ratio Performance figures of column internals, technical data Liquid load, vapor load (loading factor, operating range) Partial load behavior, separation efficiency (no. of theoretical separation stages per m) Stage specific pressure drop Required specific column volume (column volume/vapor flow rate) Costs Column costs, peripheral equipment costs, operating costs Reliability

0

Influence by process or system*

0

0

0

Main resistance of transfer process in liquid phase in gas or vapor phase Discharge of intermediate fractions required Constant temperature along tower required Foaming tendency Clogging tendency (suspending solid) Diswetting tendency Corrosive medium Viscous liquid Heat exchange required Hydrodynamic properties Low liquid holdup allowed Low pressure drop allowed Wide range of gas and liquid rates Large gas loading Very small liquid rate Equipment demands Frequent cleaning required Low weight allowed Low construction height Small base area 0 less suitable

x very suitable

* Representation according to REICHELT [2.109].

Tray column

X

0

Trickle packed tower apparatus 0 X

X

0

0

X

0

X

X

0 0

X

0 0 (XI

X X

Tower with packed tubes

0

X

0

X

X

0

X

0

(XI

X X

trickle packing tower 0 0

0

X

X

0

217

218

2 Distillation and Partial Condensation

nals and the commitment of the operation conditions results from an economic compromise. The main criteria for rectification units that are looked for are minimum total costs, safe operation and no environmental pollution. If the investment costs of column internals are to be compared, the solution of a separation problem is first analyzed. Suitable internals have to be specified with criteria for comparison purposes, such as the specific column volume v, as the column volume per vapor or gas load, and separation efficiency for the optimum operating condition. For random and regular packing, 71

&-.Z 4

Z

HETS

(2-158) For trays with a constant tray spacing Az, 71

d’*--N,* A z 4

Performance figures and cost data are given in Fig. 2-80 and Table 2-32. Economical optimization of a rectification unit regarding the reflux ratio is discussed in Chapter 2.5.2.5, Fig. 2-52. Further optimization, especially with a view to a heat network are presented by BILLET[2.2, 2.141. Several opportunities exist for process control of rectification units, depending on the requirements [2.2, 2.130-2.132, 2.1431. Figure 2-81 shows a simplified section of a flow sheet for a rectification unit with the usual process control equipment.

2.7 Rectification Unit Accessories In addition to the internals (trays, random packing, regular packing) required to increase the phase surface area, additional elements are provided in the column [2.2]: 0

0 0

AZ

0

(2-159)

T~ obtain the specific column costs c for comparison of different types of internals, only has to be multiplied with the interrials specific costs, C, per m3 column volume. c, has to be determined for each thosen column diameter d

c = v.c,

(2-160)

A comparison and evaluation of different column internals is given by BILLETin [2.2].

Support plates for filling material (Fig. 2-82) and support grid for packings Hold-down plates (Fig. 2-82) Liquid distributor for filling material and packings (Table 2-33) and liquid redistributors (Fig. 2-82) Demister between rectification trays and in front of vapor lines (wire gauze or plastic gauze with a strengthened cover [2.134], etc.)

Tube evaporators with thermosiphon circulation or continuous flow evaporators are mainly used in rectification facilities. Falling film evaporators or film evaporators are seldom used. The vapor is condensed in a horizontal liquid-cooled tube bundle heat exchanger or air-cooled finned-tube heat exchangers [2.135].

2.7 Rectification Unit Accessories

219

L \

..-

I " I

0

Carbon steel column shell

I

I

I

1

2

Tower diameter Iml-

I I

3

Fig. 2-80. Relative column costs related to the empty column as a function of the column diameter. Representation according to REICHELT[2.109]. Column data: d = 1 m, I/= 1 m3 -. -. Pall rings St. 50 x 50 - Bubble cap tray - .. . - Pall rings St. 25 x 25 Centrifugal tray Pall rings St. 15 x 15 Channel tray - _ _ Interpack 20 x 20 ..... Valve tray 1 m3 packing with shell 1 m 3 column volume including trays and shell, tray spacing 0.5 m

Table 2-32. Performance figures and pressure drop [OJ], [2.132]

*,

Column internals

F-factor

Number of theoretical

Specific tray pressure drop Ap/N, (N/m2)

Kittel centrifugal expanded-metal tray Valve tray Sieve tray Channel tray Bubble cap tray Grid packing, coarse Metal gauze packing Grid packing, fine Pall ring 25 x 25 x 0.6 Pall ring 5 0 ~ 5 0 ~ 1 . 0 Interpack filling 20 x 20 x 0.4

2.89 2.22 2.12 1.74 1.54 3.04 2.82 2.59 2.30 2.47 2.18

0.65 1.43 1.40 1.52 1.34 0.75 3.69 0.71 3.02 2.18 2.08

558 535 576 936 694 171 140 414 242 238 46 1

*

Values are obtained with the system ethyl benzene - styrene at 133.3 mbar, at 80% of the flooding point, and a reflux ratio (v = 03). Tray spacing Az = 0.5 m.

220

2 Distillation and Partial Condensation

HE1

d

Fig. 2-81. Simplified process flow diagram of a rectification unit. C 1 Column TIC 1 Overhead product, temperature H E 1 Reboiler control FIC 1 Feed control HE 2 Condenser Top product cooler HE 3 FR 1 Top product, flow, recorded Bottom product cooler, feed Bottom product, flow, recorded HE 4 FR 2 preheater Bottom product, temperature, TR 2 recorded V 1 Top product receiver PIC 1 Steam pressure indicator and control, Process control keys: P Pressure possible response to total column L Level pressure drop F Flow PdI 1 Pressure drop, indication T Temperature PI 1 Top column pressure I Indication Bottom column pressure PI 2 LIC 1 Level control, bottom R Recording TR 1 C Control Overhead product temperature, recorded FRRC 1 Reflux ratio control, possible response to top product temperature

2.7 Rectification Unit Accessories

221

1. Support plates*

a) Corrugated expanded grid or perforated plate welded on a frame, diameter up to 1000 mm and a low load (one piece)

b) Frame with flat profile grid for small and intermediate diameters, and intermediate load, with and without expanded grid (also used as hold-down tray)

c) Braun support grid, one piece, for diameters up to ca. 600 mm made from standard parts (500x 250), edge segments are adjusted, suitable for every diameter, grid height up to 150 mm

d) Standard support tray made from single beams, for diameters larger ca. 950 mm, light and heavy model

2. Hold-down plates** (bed-limiter)

e) Sieve or expanded metal plate for a diameter up to ca. 1200 mm Fig. 2-82. Packed tower internals.

0 Sieve or expanded metal including weight for diameters larger ca. 1100 mm

222

2 Distillation and Partial Condensation

3. Liquid redistributors**

g)

Rosette redistributor for diameters up to

600 mm

h) Redistributor for diameters larger ca. 1200 mm

Fig. 2-82. (continued)

* Representation according to Vereinigte Full-

korper-Fabriken GmbH & Co., RansbachBaumbach. ** Representation according to Norton GmbH, Wesseling.

2.8 Parallel Flow Distillation With parallel flow distillation [Oh] vapor and liquid are guided in parallel flow through a column in contact zones, mainly tube bundles (Fig. 2-83). Upflowing vapor drags liquid upward on the inner tube wall in the form of a film, where mass and heat transfer occur in a similar manner to rectification. In general, during phase contact, vapor is enriched with the low-boiling component, while the liquid phase is reduced. At the respective ends, phases are separated by a separator; vapor enters the next higher tube bundle, while liquid flows to the next lower tube bundle. Vapor arriving at the top of the column is condensed in a total condenser. The condensate is divided into the overhead product and the reflux which is fed back to the column. Some of the bottom product is discharged, and the rest is fed back to the distillation still. With counterflow distillation, vapor velocities are 1-3 m/s, whereas in parallel

flow distillation considerably higher gas velocities, up to 50 m/s, are possible. The separation efficiency decreases at increasing allowable vapor velocity, with increasing diameter of the individual tubes in a tube bundle. More intensive phase contact with developed two phase flow is reached by vortex flow or devices to reinforce spray vortex flow, compared with pure liquid flux. Internals for counterflow columns in distillation and absorption are currently being developed. The use of this type of column on an industrial scale is not yet known.

2.9 Nonadiabatic Rectification Nonadiabatic rectification (thermal rectification, redistillation) is a thermally gentle separation process of high-boiling liquid mixtures, using a combination of partial distillation and partial condensation. The fractionator is a special thin-layer evaporator, with externally heated columns and an externally cooled rotor (Fig. 2-84). The vapor mixture, generated in a thinlayer or falling film evaporator, flows countercurrently in the fractionator, with the downward liquid reflux at the heated inner wall of the evaporator. By repeatedly vaporizing and condensing some of the

2.9 Nonadiabatic Rectification

4

4

223

downflowing liquid and the upflowing vapor, the low-boiling mixture components are enriched in the vapor, while the higherboiling components are enriched in the reflux. Vapor at the top of the fractionator condenses in a condenser. Some of the condensate is withdrawn as overhead product, the remaining forms the reflux which returns to the fractionator. Reflux leaving the sump of the fractionator flows back to the distillation unit and is there withdrawn as the bottom product. Depending on the separation task, the fractionator may be operated only as an enrichment column, only as a stripping column or in a normal operating mode as shown in Fig. 2-84. The thin-layer evaporator is most effective when the ratio of the mass flows vaporized and condensed on the side is equal to 1 [2.136]. Therefore, the separation efficiency of the fractionator depends on how often the generated vapor mixture is condensed and revaporized in the unit. At infinite reflux ratio and at constant relative volatility a, the separation efficiency in the concentration range of interest is Vapor phase Liquid phase

Fig. 2-83. Two stage cocurrent tube bundle distillation unit. ST Still TC Tube bundle column C Condenser TP Top product RF Reflux BP Bottom product SE Separator EA Entrainment area

(2-161) where the vapor concentration at the top and bottom of the fractionator n, = N,/Z Number of theoretical stages per m of the fractionator z height for mass and heat transfer of the fractionator (chosen upper limit of the dimensions of a thin-layer evaporator : 850 mm diameter, 9 m height)

yo, yL?

224

2 Distillation and Partial Condensation

Table 2-33. Liquid distributors. Presented by MANTEUFEL [2.133]. System

Method

1

Construction

Drain-

Liquid flux per drain

Om

No. of drains per m2

Liquid rate (m3/(m2 h))

> 10

40-400

Medium >1

40-1000

Medium to very large

+

liquid jet # '3

,

3'6

.

Hole

1. Hole or sieve tray distributor with or without vapor riser (tube) 2. Channel or trough-type distributor Tube distributor (straight or concentric)

like liquid

>5

1. Spout or riser distributor 2. Channel or trough-type distributor

2.9 Nonadiabatic Rectification

Operating range

Energy

con-

sumption

Free vapor area

Functional range

Column diameter (m)

Performance

Pros

Cons

225

Limited 51

No

20-40% > 50%

Normal pressure overpressure

Up to 2.5 2 0.6

Sensitive to dirtying or fouling

Equal single drain fluxes

Few drains

Unlimited

No

20-40% > 50%

Normal pressure overpressure

Up to 2 20.6

Sensitive to dirtying or fouling

Wide operating range

Not for small liquid rates

(continued next page)

226

2 Distillation and Partial Condensation

Table 2-33. (continued) System

Method

Atomization to droplets or fog

1

t! I

' 111

\\

I/ 11\ I/ 11 1I 11

Construction

Liquid flux per drain Wh)

No. of drains per m2

Liquid rate (m3/(m2 . h))

-

M

Small to very large > 0.2

30-1600

\\\

1 1 1

Nozzle

1. Solid-cone nozzle

Flat-spray nozzle Slot nozzle Perforated die (sprinkler) Beam nozzle Deflector nozzle Two-fluid nozzle 2. Pipe distributor with nozzles Perforated pipe distributor

$1 6o

20 - 5000 Mechanical A drive

+' Effective side pressure

Roloiion

Atomization to droplets or fog Liquid jet or droplet discharging

-

> 100

a63 Rotary disk

Segner sprinkler, single or several arms

03

<40 rotating

Medium to large 0.3-2 >3

2.9 Nonadiabatic Rectification

227

Operating range

Energy consumption

Free vapor area

Functional range

Column diameter (m)

Performance

Pros

Cons

Limited
Yes

Almost 100%

Vacuum

Unlimited

Very sensitive to dirtying or fouling possibility of clogging

Large distribution effect

Unequal sprinkling and fog forming

Limited <1:4

Yes, Almost but no 100% external mechanical drive

Normal to overpressure

Up to 2

Sensitive to dirtying or fouling

Large distribution effect

Unequal springkling

Up to 1

(continued next page)

228

2 Distillation and Partial Condensation

Table 2-33. (continued) System

Method

Construction

Liquid flux per drain ( W

No. of drains per m2

Liquid rate (m3/(m2 . h))

0.05-0.5

Up to 18000

Very low to medium 0.5- 10

>2

2001000

Very low to high

5

B

Trickling filter

Discharge head or splash plate Trickling filter 6

Siphoning effectby capillary film

,+,,!*#,,* l I , 1 , 1 1 1 1

x x T x T-

.;x'\*k%,

Capillory

Capillary tray

7

Liquid jet pass

Profiled slot

Profiled slot distributor

> 0.5

2.9 Nonadiabatic Rectification

229

Operating range

Energy consumption

Free vapor area

Functional range

Column diameter (rn)

Performance

Pros

Cons

Limited

No

Ca.80 070

Normal pressure to overpressure

Up to 1

Sensitive to dirtying or fouling

Simple construction

Unequal sprinkling

Wide up to 1 : 10

No

50-70070 Fine

Unlimited

Very sensitive to dirtying or fouling

High no. of drains

Very sensitive to dirtying or fouling

Very wide up to

No

50-60%

Unlimited

Sensitive to dirtying or fouling

Small and equal single drain fluxes

Small manufacturing tolerances

<1:3

1 : 10

vacuum

Fine vacuum to overpressure

230

2 Distillation and Partial Condensation

The advantage of rectification in a thinlayer evaporator in the pressure range of 1-25 mbar is the small operating volume of the rectificator and the distillation unit. From this, a lower mixture residence time, a small specific pressure drop of a transfer unit, high vapor load and a low sensitivity to fouling mixture components results.

2.10 Partial Condensation TP

View A-A

According to the discussion in Chapter 2.2, liquid mixtures containing different boiling components are partially separated by supplying heat to the system. Analogous to this, vapor mixtures with different condensing components are partially separated by partial condensation which is achieved by removing heat from the system. For example, for a vapor mixture D,, with a mol fraction y , (of the lower-boiling key component), leaving an exothermic reaction cooled by vaporization, or a continuous flow evaporator, an amount R is separated by partial condensation (Fig. 2-85). The remaining vapor DD = bF- R , leaving the partial condenser is then enriched with the lower-boiling mixture components, which gives yD>y,. The enrichment

Fig. 2-84. Thermal rectification fractionator with thin-layer evaporator as reboiler. Representation according to Buss - SMS Chem. Eng., Zurich/Butzbach. H C Heated column CR Cooled rotor C Condenser TE Thin-layer evaporator VA Free area for upflowing vapors F Feed R Reflux TP Top product BP Bottom product

2.10 Partial Condensation

2

\

/

&YD

231

The dephlegmator arrangement is shown in Fig. 2-86. According to Fig. 2-44 the enrichment line in the McCabe Thiele diagram fixed by the coordinates x, and y,, with the starting point A is necessary to determine the number of theoretical separation stages. Since yo > y, the dephlegmator acts as a separation stage. In Fig. 2-88 dL is the condensed vapor in the height element dz of the partial condenser of total height Z . A mass balance of a differential element dz for the condensed amount of vapor di gives

(2-163)

Fig. 2-85. Vertically mounted partial condenser. 1 Vapor inlet 2 Vapor outlet 3 Condensate outlet 4 Cooling agent inlet 5 Cooling agent outlet

yD

Upon integration, with y between the limits yF and yo and L between 0 and R, respectively, gives

(2-164)

- y,

increases with increasing condensate amount. If a partial condenser is linked to a rectification column, as much vapor condenses as is needed for the reflux of the column. A partial condenser is then named a “dephlegmator”, as shown by Fig. 2-86. The enrichment yD - yF in the dephlegmator increases with increasing reflux ratio v = R/DD,shown qualitatively in Fig. 2-87. A mass balance over the partial condenser gives the molar fraction x, of the key component in the condensate

(2-162)

where x is the concentration of the condensate at the contact area with the vapor and y is the corresponding vapor concentration at any horizontal reference cross section of the partial condenser. Neglecting the mass transfer resistance in the vapor, y is approximately the phase equilibrium concentration corresponding to x. Therefore, Eq. (2-164) can be evaluated analogously to Eq. (2-3) or (2-14). For example, for the case of constant relative volatility a, in the concentration range of interest, y, Iy Iyo, it follows from Eq. (2-164) that

=

k)A

. I-.)-( 1-Y, -YO

1 ~

-1

(2-165)

2 Distillation and Partial Condensation

232

, i A

RC L-

.

I

Fig. 2-86. Partial condenser and rectification.

RC Rectifying column PC Partial condenser C Condenser cooler

Fig. 2-88. Differential section of a partial condensation Eq. (2-163).

CA Cooling agent CO Tube wall and heat exchanging circurnference of a partial condenser FF Falling film VP Vapor phase

t

0

A

YF

-

Fig. 2-87. Enrichment of the lower boiling cornponent by partial condensation as a function of the reflux ratio v (qualitatively). yo Remaining vapor fraction of the low-boiling

component

yF Low-boiling component vapor fraction in the

The separation efficiency of a partial condenser is usually less than the possible theoretical efficiency, which is distinguished by the phase equilibrium between the vapor bulk and the condensate surface. The difference between the true and theoretical separation effect is mainly due to the condensate load and mass transfer. Calculation methods are presented in [2.137], which enable estimation of the separation efficiency, depending upon the operation conditions, apparatus dimensions and the mixture properties. The height of a partial condenser 2 for mass and heat transfer is given by

-

2.10 Partial Condensation

The amount of condensate R based on Eq. (2-164), is

R = D F . [I - exp

(-7 A)]

-

,c

G!iC

c F

(2-167)

233

Y F Y - x

where

Ah,,g mean vaporization enthalpy along 2 k

e

d At9 dn dF

I

I

/cc

related to the mean temperature and concentration mean heat transfer coefficient number of tubes in the partial condenser diameter of a single tube mean temperature gradient between t9D and VF mean condensate temperature along- Z mean temperature of the cooling medium along 2

Partial condensation may also be carried out as partial countercurrent direct condensation in columns [2.2, 2.138, 2.1391. For example, in a reaction vessel using evaporative cooling, the mixture vapor generated by the reaction flows in a controlled countercurrent manner with the cold feed to the reaction vessel (Fig. 2-89). The cold feed is preheated by partially condensing the mixture vapor. During the phase contact, the condensate formed flows back to the reaction vessel with the preheated feed, the raw material, or reactants. An advantage of direct condensation of the mixture vapor and cold feed in countercurrent flow is that only a small portion of the mixture vapor has to be condensed in the dephlegmator of the column. In addition, lower-boiling compo-

Fig. 2-89. Chemical reactor with countercurrent

direct condensation column. CC Countercurrent direct condensation column R Chemical reactor C Condenser CF Cold raw material feed PF Preheated raw material V Vapor, generated by chemical reaction and/ or by evaporative cooling RC Remaining condensate

nents of the reaction mass can be economically continuously withdrawn. In the field of environment protection, partial condensation is applied to separate and recover condensable solvents from those gases which are noncondensable under the normal operating conditions.

234

2 Distillation and Partial Condensation

References [2. I] [2.2] [2.3]

[2.4]

[2.5] [2.6] [2.7] [2.8] [2.9] [2.10]

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235

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236

2 Distillation and Partial Condensation

ERBAR,R. C., JOYNER, R. S., and MADDOX,R. N.: Petro/Chem. Eng. 33 (1961) 3, c19-c 22. WINN, F. W.: Pet. Refiner 40 (1961) 4, 153-155. [2.65] VOGELPOHL, A,: Chem. Ing. Tech. 46 (1974) 5 , 195. [2.66] NEUMANN,K.-K.: Erdoel Kohle 26 (1973) 4, 198-202. [2.67] FUTTERER,E., LANG, G., and NEUMANN,K.-K. : DECHEMA Monogr. No. 1410-1431 73 (1974). [2.68] KETCHUM,R. G., and JAKOB, C.-H.: krfahrenstechnik 15 (1981) 8, 548-552. [2.69] BILLET,R. : Optimierung in der Rektifiziertechnik. Bibliographisches Institut, Mannheim 1967. [2.70] KLAPP,E. : Apparate- und Anlagentechnik. Springer-Verlag Berlin, Heidelberg 1980. [2.71] BTZE,H.: EIemente des Apparatebaus. Springer-Verlag Berlin, Heidelberg 1967. [2.72] JORDAN,D. G. : Chemical Process Development, 2 Vols. Interscience Publishers, J. Wiley & Sons, New York 1968. J. : Grundlagen der Dimen[2.73] STICHLMALR, sionierung des Gas/Fliissigkeit-Kontaktapparates Bodenkolonne. Verlag Chemie, Weinheim 1978. [2.74] HOPPE, K., and MITTELSTRASS, M.: Grundlagen derDimensionierung von Kolonnenboden. Steinkopff, Dresden 1967. [2.75] FOLDES,P., and EVANGELIDI,I.: Br. Chem. Eng. 13 (1968) 9, 1291-1293. [2.76] SUNDERMA", U.: Chem. Tech. 20 (1968) 3, 145-150. [2.77] CHASE, D. J.: Chem. Eng. 7 (1967) 105-116. [2.78] NERETNICKS,I.: Br. Chem. Eng. 15 (1970) 2, 193-198. [2.79] BRAUER,H., and WIELE, H.: Chem. Zng. Tech. 45 (1973) 7, 481-484. [2.80] AIChE.: Bubble Cup Design Manual. American Institute of Chemical Engineers, New York 1958. [2.81] DIERY,W.: Dissertation, TH Munchen 1962. , Chem. [2.82] BRAUER,H., and T ~ I E L EH.: Zng. Tech. 45 (1973) 9/10, 617-620. [2.83] Company report: ,,Im Dienste der chemischen Industrie," Koln-Niehl.

ACV Dr. Stage, Schmidding und Heckmann. [2.84] Company report: ,,Varioflex-Boden fur den Stoffaustausch," Stahl GmbH Viernheim. [2.85] Company report: ,,Veroffentlichungen uber Thormann- und Streuber-Beden," Montz GmbH Hilden. [2.86] RAICHLE,L., and BILLET,R.: Chem. Ing. Tech. 35 (1963) 12, 831-836. [2.87] RODL, R.: CZ Chem. Tech. 3 (1974) 3, 93-97. [2.88] Company report: ,,GHH-Destillationselemente," Gutehoffnungshutte AG Oberhausen-Sterkrade. [2.90] Company report: ,,Informationsschrift," Montz GmbH Hilden. [2.91] DIERY,W.: Chem. Ing. Tech. 47 (1975) 23, 964-970. [2.92] MOLZAHN,M., and SCHMIDT,R.: Verfahrenstechnik 9 (1975) 8, 388-395. [2.93] SOUDERS,M., and BROWN, G.: Ind. Eng. Chem. 26 (1934), 98-103. J. : Dissertation, TU Mun[2.94] STICHLMAIR, chen 1971. [2.95] SATTLER,K. : Thermische Trennverfahren. Aufgaben, Losungen, Auslegungsbeispiele. Vogel-Verlag, Wiirzburg 1979. [2.96] From [2.2]: After calculations from Fa. Koch Engineering, Wichita, Kansas/USA and Koch International, Bergamo, Italien. [2.97] STICHLMAIR, J. : Dissertation, TU Munchen 1971. [2.98] LEWIS,W. : Ind. Eng. Chem. 28 (1936) 4, 399-402. [2.99] STICHLMAIR,J., and WEISSHUN,E. : Chem. Ing. Tech. 45 (1973) 5, 242-247. [2.100] STICHLMAIR, J. : Habilitationsschrift, TU Munchen 1978. [2.101] COLBURN,A.: Znd. Eng. Chem. 28 (1936) 5 , 526-530. [2.102] ZELFEL,E.: Chem. Zng. Tech. 40 (1968) 7, 327-332. [2.103] MUHLE, J.: Chem. Ing. Tech. 44 (1972) 1/2, 72-79. [2.104] KUTZER,H.: Chem. Ing. Tech. 42 (1970) 3, 128. [2.105] HOPPE,K.: Wiss. Z. Tech. Hochsch. Magdebzrrg 11 (1967) 3/4, 493-501. [2.106] NITSCHKE,K., and OPITZ, H.: Chem. Tech. 20 (1968) 1, 23-25.

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Thermal Separation Processes: Principles and Design Klaus Sattler, Hans Jacob Feindt copyright 0VCH

Verlagsgesellschaft m h H , 1995

3 Absorption

3.1 Principle of Absorption and Desorption, Processes and Process Examples 3.1.1 Concepts and Process Examples Absorption is the acceptance and dissolving of vapors and gases in liquids [3.1-3.12, 3.761. Absorption is a thermal separation process using an auxiliary substance, wash liquid, solvent, or absorption liquid. Gas or vapor mixtures may be partially separated using absorption. During contact of the gas mixture and the solvent in an absorber, one or more components which are of low concentration in the original mixture, are dissolved as selectively as possible by the solvent. The gas mixture becomes reduced in these components, and a separation occurs. Depending on the type of gadsolvent absorption system, the result is a physical solution (physical gas scrubbing, physisorption) or a chemical compound (chemical gas scrubbing, chemisorption) consisting of gas and solvent. The component exchanged between the feed phase and the absorption phase is called, in the nonbonded state, the absorbate, also key component, and in the bonded state, the absorbed component; the receiving wash liquid is called the absorbent, and the loaded solvent phase is the loaded absorbent (solution). Absorption processes are mainly used to separate gas mixtures and to purify gases. With a gas separation, the dissolved com-

ponents are the desired products, and with gas purification, unwanted pollutants are absorbed. In the latter case absorption is termed “wet gas scrubbing”, sometimes combined with wet dust removal and is, therefore, an important pollutant removal process in the field of environmental protection [3.13]. The loaded solvent is usually regenerated by reversing the process, by desorption or stripping. After the solvent is recovered, it is ready for reuse. The solvent is regenerated in a desorber linked downstream to the absorber (Fig. 3-1), and the recovered solvent is hence recycled to the absorption unit. Since gas absorption is favored under low temperature and high pressure (see Chapter 1.4.3.3), desorption is carried out under high temperature and low pressure. With chemical absorption, the reaction rate increases with increasing temperature. Chemically active wash liquids are characterized by high selectivities and good load capacities even under low absorbate concentrations in the gas phase. The absorption process sometimes requires several absorbers connected in series. If the same solvent is used in the different stages, the operation conditions have to be decisively changed from stage to stage to obtain a good absorption yield. The operation temperature has to be lowered from stage to stage and the pressure has to be increased. More commonly different types of solvent are employed, in which a solvent with higher a selectivity and load capacity, compared with the previous stage, is used in the stage. A combination of physisorption as a crude pre-step

240

3 Absorption

UE2 -I

AS

PS

CG ---I

RS 4

Y HE^ Fig. 3-1. Gas absorption unit with solvent recovery. AS Absorber DS Desorber (recovery column) SE Separator H E 1 Recovered solvent cooler, first stage HE 2 Recovered solvent cooler, second stage C Condenser R Reboiler CG Crude gas PG Purified gas SA Solvent addition RS Rich solvent PS Recovered solvent EG Exhaust gas

and chemisorption for the removal of the remaining absorbate is often applied.

3.1.2 Process Examples Some examples for the separation of gas mixtures and for gas cleaning by absorption processes are listed in Table 3-1. Fig. 3-2

shows a schematic for a two stage absorption process of WIEGANDto produce 15% hydrochloric acid from a waste gas containing hydrogen chloride. The absorption process is explained in the flow diagram (Fig. 3-2b). A simplified flow sheet for the desulphurization of flue gas by the WALTHER process is presented in Fig. 3-3: Ammonia is added to dust free, hot flue gas on the pressurized side of the blower. After passing through a regenerative heat exchanger (W4) the flue gadammonia mixture is treated in the first wash tower (B4) with a wash solution, formed by the ammonia, acidic components of the flue gas and water fed into the tower. Here the following reactions proceed simultaneously : HC1+ NH3 H F + NH, SO2 + NH3 + H2O NH4HSO3 + NH3 (NHJ2SO3 + 1/2 0

+ N&Cl =. N&F

+ NH4HS03

*

2

*

(NH4)2S03 (NH&S04

The salts formed remain in the wash solution, in which the salt accumulates close to the saturation limit. The considerably cleaned flue gas then passes to the mist eliminator (F4) and into the second wash tower (B5),to be treated with fresh water in order to remove most of Fig. 3-2. Absorption of hydrogen chloride with b water as solvent in two stages. Representation according to GEA Wiegand GmbH. a) Absorption unit, schematic 1 Jet scrubber (1st stage) 2 Jet scrubber (2nd stage) A Crude gas B Purified gas C Fresh water D Hydrochloric acid b) Absorption process in an operating diagram X HCl ratio in water Y HC1 ratio in air

_I D

B

0

?-

I

b)

Crude gas Acid

0.0°6

n


0

0

~~

1. Stage

2.Stage

> Purified gas Water

1

I

t

Y

X [kmol/kmoll

-

1

3.2 Requirements of the Wash Liquid or Solvent, Solvent Consumption

the carried salt solution and the remaining ammonia, In the heat exchanger (W4), the flue gas is heated, finally leaving the plant via a stack. The salt solution from the first washer of the wash cycle, containing mainly ammonium sulfite, oxidizes to ammonium sulfate in the oxidizer (B2). In the spray dryer (B3) the salt is dried to obtain a crystalline dust consisting of ammonia sulfate with chlorine and fluoride components, by using some of the dust-free flue gas at a temperature of 350-400°C. The crystalline dust is separated by an electrostatic filter (F3) and is collected in product storage facilities (Bll). The next step is the pelletizing of the crystalline dust in a balling pan (PTl) using water as a binding agent, followed by drying in a drum dryer (Tl), again using some of the dust-free flue gas. By means of a screen fractionator (SI), the desired particle sizes fall into the product silo (BS), oversized particles are guided to the mill (MI) and undersized particles are recycled to (B11). The advantages of this process are: generation of a granular nitrogen fertilizer in a directly usable form without generation of wastewater, regenerative reheating of the treated flue gas, and the relative insensitivity to crusting since only water soluble salts are formed. With additional separation of nitric oxide (NO,) from flue gas, a third wash stage is required. The desulfurized flue gas is

243

mixed with air containing ozone, and nitrogen monoxide is thereby oxidized to nitrogen dioxide which is water soluble. The ammonia water solution absorbs the generated nitrogen dioxide and a subsequent oxidation yields ammonium nitrate.

3.2 Requirements of the Wash Liquid or Solvent, Solvent Consumption The solvent must meet certain requirements to ensure an economic, safe and environmentally friendly absorption process. These requirements also necessary for selection of the solvent are listed in Table 3-2. At least three components are present in the absorption process; the inert carrier gas, the solvent as an auxiliary substance, and the gas component i to be absorbed. Under steady-state operation, flow rates of the pure solvent i, and the inert carrier gas GTare constant. The amount of solvent required assuming a low vapor pressure of solvent and neglecting solubility of the carrier gas in the solvent, follows from a mass balance over the absorber (see Fig. 3-4) (3-1)

Fig. 3-3. Desulfurizing of flue gas, Walther process, simplified flow sheet. Representation according to FISCHER[3.14]. Example: Flue gas treatment in the Mannheim main power plant. Process data according to [3.15]: Flue gas input in FTP: 753000 m3N/h Input gas temperature in FTP: 130°C Input SO, concentration: max. 2200 mg/m3 Input dust contents: 50 mg/m3, Remaining SO, concentration after heating before stack: <200 mg/m& Ammonia consumption: 0.8 t/h at max. sulfur concentration Amount ammonium sulfate: 3.2 t/h at max. sulfur concentration FTP: Flue gas treatment plant

. Examples for the separation of gas mixtures and for gas cleaning by absorption [3.9, 3.17-3.213. Absorbate ______~~

Absorbent

Reaction equation or absorption produ

___

rption HCI rption SO, bsorption solvent HZS, CO,, COS, CS2

bing

water HZS04,aq

water, wash oil polyethylenglycol dimethyl ether

HCI,, concentrated sulfuric acid solvent H2S, CO,, CS,

H,S, CO,, NH,, COX, HCN methanol H,S, CO,, NH,, HCN CO2, HzS N-methylpyrrolidone C02, H2S tetrahydrothiophendioxide(sulfolan)/ H,S, CO, H2S, COZ, COS diisopropanolamine/water soda lye ( ~ V O )cold ,

co,, H2S

H2S

ubbing

CO,

ubbing

2 NaOH + CO, 4Na2C0, + H,O 2NaOH + H2S Na2S + 2 H,O 4NaOH + COS Na2C0, + Na2S + 2 KZCO, + CO, + H,O + 2KHC0, +

cos

rubbing C02 LD, .. . etc.) COz, H2S

SO,

ubbing

SO2 process RBERGKOBE , . etc.) SO2 n-Lord

soda lye (2-4%), hot aqueous K,CO, solution (10- 12%), cold aqueous K,CO, solution (15-30%), hot aqueous monoethanolamine mixture (10-20%) aqueous diethanolamine mixture (10-25 070) aqueous Ca(OH), or CaCO, solution, if necessary + additive aqueous Ca(OH), or CaC0, solution aqueous Na2S03 solution

SO2, NO, process a scrubbing C 0 7 crubbing CO,, H,S

coe process

HZS

ammonia water ammonia water alkazid M acid (aqueous potassium salt solution of methyl amino propionic acid) Na3As04 solution

+

K2C03+ H2S+KHC03 + KHS 2HOC2H4NH, (HOC,H,),NH Ca(OH), Ca(OH),

+ SO,

+ COz + H 2 0 + H2S

+

+

+

-+

(HOC

(HOC2H4)2N

CaSO,

+ SO, + 1/202

Na2S03+ SO2 + H 2 0

+ H,O +

CaSO,

+H

2NaHS0,

S02NH, + H 2 0 4NH,HS03 etc. 2NH3 + CO, + H,O + (NH,),CO,

+ H,S+

H3C H3C>N-CH-COOK

H3C>N - CH - C H3C Na,AsO, + H,S + Na3As03S+ H 2 0 A Na3As0,S + 1/2 0, Na,AsO, + S Re -+

-

3.2 Requirements of the Wash Liquid or Solvent, Solvent Consumption

245

Table 3-2. Requirements on solvent and wash solution. Selection aspects. 0 0

0

0

0 0 0 0 0 0 0

0

Sufficiently good solubility of the key (gas to be absorbed) component (favorable absorption equilibrium properties). High selectivity to the key component (solvent acting physically (physisorption) maximum selectivity is 5 : 1, i. e., at least 1/5 of the absorbed molecules are of other components of the gas mixture than the key component. In chemically acting solvents selectivity is considerably higher). Simple to regenerate (no azeotrope formation between solvent and absorbed component, no formation of chemical compounds which are not or only costly to regenerate, high boiling point difference between solvent and absorbed component, favorable desorption equilibrium properties). Low vapor pressure at absorption temperature to avoid solvent loss by evaporation and, therefore, impurities of the cleaned gas flow by solvent vapor. Moderate boiling point to separate the dissolved component by partial solvent evaporation. Low melting point to tolerate a high temperature gradient between absorber top and bottom. Large specific heat to buffer absorption heat due to solvent heating. Low viscosity. Chemical and thermal stability to avoid corrosion and solvent decomposition. Good availability, low price. If possible the solvent should not be toxic or should not have any environmental impact, the solvent flash point should be high, and it should not have a tendency to foam by itself or with elutriated gas components. Low solvent residue should not have an impact on the absorption product value.

where Y,, Y, concentration of the key component i in the gas phase at the entry and exit of the absorber (kmol i/kmol inert carrier) X , , X , concentration of the key component in the solvent (liquid) at the entry and exit of the absorber (kmol i/kmol liquid phase solvent)

(The use of the molar loads instead of mole fraction offers an advantage in calculations. In particular the balance line Y(X) at GT = const. and L , = const. are straight lines in the Y(X)-loading diagram, which will be shown later.) The molar flow rates GT and L , valid in Eq. (3-1) are

(3-2) and L,=--

LCl

1

+xa- L,

*

(1 - x u )

(3-3)

where G, and L , are the entry flow rates of the gas mixture and solvent, and xu and y, are the absorbed component and absorbent entry mole fractions of the liquid and gas phase, respectively. If evaporation of the solvent and solubility of the inert carrier has to be considered, the total flow rates L and G, and the mole fractions x, y are used. At least three bal-

Y h

i

/

4 I

Fig. 3-4. Material and enthalpy balances of cocurrent (a) and countercurrent (b) absorbers.

ance equations are obtained describing the mass balance over the absorber. It is sometimes practical, to recycle some of the loaded solvent, with intermediate cooling through a section of the absorber, or to recycle through the entire absorber (Fig. 3-5). Therefore, the requirement for fresh solvent is (3-4)

Xa.1

=

Xa+r-Xw r+ 1

(3-6)

An increase in the recycled solvent flow rate L,, and hence the recycle ratio r, implies a saving in wash solution, but also a decrease in the mean driving force for the transfer of absorbate from the gas into the liquid phase (see Fig. 3-7). The result is a necessary increase in the interphase surface for mass transfer.

and

r=7

LTU LT

(3-5)

where r is the recycle ratio between the flow rate L,, of the recycled solvent and the fresh or regenerated wash liquid L,. A mass balance over the mixing stage (Fig. 3-5) for the key component gives an absorbed component loading X u , , for the wash solution fed to the absorber

3.3 Enthalpy and Heat Balances An enthalpy or heat balance for an absorber without heat losses as Shown in Fig. 3-4 gives

.

Ga ' &,a + =

*

Zi,a f

I% ' + (Ya-Yw)

Gu * hg,@+ L , . hC,@+ Q

*

1

f i ~ b=

(3-7)

3.3 Enthalpy and Heat Balances

iT x, I

? G a s ph.ase

4

Liquid phase

Fig. AS C P

3-5. Absorber with solution recycling.

Absorber Cooler Pump

MS Mixing stage

where

hgLh, enthalpy of gas and liquid phases Ah,, Q

absorption enthalpy heat flow removed from the absorber by cooling (Q = 0 under adiabatic absorber operation)

During the absorption of gases in solvents the heat of absorption is released and absorption is therefore an exothermic process. The heat of absorption is mainly taken up by the solvent, as heat transfer within the gas flow is poor and negligible. Therefore, during an adiabatic absorption process the temperature of the solvent increases

247

due to the heat of absorption from an entry temperature V/,a to a n exit temperature V,,,. V , , may be estimated by Eq. (3-7) with the assumption, that all of the heat of absorption is taken up exclusively by the solvent. Since an increase in the temperature from 8,, to 8,, occurs, this causes a decrease in the solvent absorption capacity of the soluted substance according to Henry’s law. The absorption efficiency reduces due to an unfavorable equilibrium position at the higher temperature, (see Chapter 1.4.3.3). Consequently, isothermic absorption is aimed for in practice. The heat flux to be removed under isothermal absorption at (V1,@= LP/,) follows from Eq. (3-7). Continuous removal of the heat of absorption is very difficult. The absorber is cooled stagewise if the absorption heat is considerable. If a volatile solvent is used, some of the absorption heat is used up by vaporization of solvent. For this case, the heat flow of vaporization is added in Eq. (3-7) on the right side. (To calculate the exact temperature and concentration profile over the active absorption column height, the appropriately adapted equation system in Chapter 1.9.3 is used.) With isobaric absorption, the absorption enthaQy AEAbis the specific heat of absorption for a kmol absorbed substance. This depends on the temperature, the absorbate-solvent system and the absorbate concentration in the solution. If the absorbate is soluble as a gas, the absorption enthalpy is the solution enthalpy. This may be determined, for example, from the temperature dependency of Henry’s constant Hi

(3-8) If the absorbate condenses under the absorption conditions, the condensation en-

248

3 Absorption

thalpy is added to the mixing enthalpy. The absorption enthalpy then becomes

where y i is the activity coefficient and h/,, is the vaporization enthalpy of the absorbed component. With chemisorption the heat effect is caused by a phase change enthalpy and a reaction enthalpy. The reaction enthalpy hRis calculated by means of the formation enthalpy fiij for each reactant j which participates in the reaction

ing diagram are governed by the solvent ratio v for each absorption stage. For example, an absorbate balance, for stage AB1 (Fig. 3-6a) gives

Cocurrent absorption is usually carried out in spray or jet scrubbers and film absorbers. Cocurrent absorption is mainly used for the absorption of chemically acting wash liquids and only small residence times are needed to obtain the required absorption yield also at a large throughput.

(3-10) where vj is the stoichiometric ratio of the reactant j . A graphical method to determine the absorption enthalpy for physical systems by the aid of the enthalpy-concentration diagram is described in [3.1]. Absorption enthalpies for practical use are given in [3.16].

3.4 Cocurrent Phase Flow Absorption In absorption processes with cocurrent flow, the feed phase flows parallel with the solvent phase through a single absorber or through a series of absorbers. Figure 3-6 shows a schematic for a cocurrent two stage absorption unit. The loading diagram corresponds to the case of the leaving phases just reach equilibrium. (If phase equilibrium is not reached, which is usually the case, a curve must be used which takes into account the true exit loading of each stage X ,, e f f , , e f f , etc., instead of the equilibrium curve in Fig. 3-6.) The slopes of the operating lines in the load-

3.5 Countercurrent Phase Flow Absorption, Design of Countercurrent Flow Columns During absorption with countercurrent flow, the feed gas phase flows upward into the absorption column. The solvent phase, which is introduced at the top and withdrawn at the bottom, flows against the gas phase. Internals in the column provide stepwise (trays, spray zones, rotating discs) or continuous (random packing, regular packing, etc.) phase contact.

3.5.1 Determination of the Column Cross-Sectional Area The column cross section AQ and column diameter d result from a flow equation with the gas phase as the reference phase (3-12)

3.5 Countercurrent Phase Flow Absorption, Design of Countercurrent Flow Columns

249

b)

a)

Y2-

y,

L 4,

2

Fig. 3-6. Cocurrent two stage absorption unit. a) Absorption unit with feed of fresh or regenerated wash solution to each stage b) Absorption unit with solution reflux (recycle) c) Operating diagram to (a) d) Operating diagram to (b) AB 1, AB 2 Absorber 1 + 2 BL 1, BL 2 Balance line for the 1'' and 2nd absorption stage EC Equilibrium curve

x,

x

-

250

3 Absorption

V,,,,, is the maximum possible effective volumetric flow of the gas phase through the absorber. It is recommended to calculate the volumetric gas flow with respect to the absorber entry and exit cross-sectional area. In that way, losses of absorbate from the gas phase, solvent evaporation, pressure drop, and increase in the temperature under adiabatic absorption are considered. The maximum allowable velocity of the gas phase, referred to the free absorber ,,, on the floodcross section w ~ , ~depends ing point or upper loading limit for the selected internals, for example, as calculated by the methods presented in Chapter 2.

units (“HTU-NTU concept”) according to Chapter 1.9. Basic equations are listed in Table 3-3. The required number of theoretical stages Nt for a countercurrent flow absorber can be graphically obtained by drawing the stages on to the operating diagram between the balance or operating line and the equilibrium curve (Fig. 3-7). From an absorbate balance over the top section of the absorption column, according Fig. 3-4, the equation of the balance line is

3.5.2 Determination of the Number of Stages and Column Height for Mass and Heat Transfer

The balance line is linear if the solvent ratio v = LT/GT in the concentration range of interest Y, > Y > Y, is constant. Solvent vaporization and carrier gas solubility are disregarded. The course of the equilibrium curve is fixed by the relationship given in Chapter 1.4.3.3

The column height Z for mass and heat transfer is determined using the concept of theoretical stages or the concept of transfer

(3-13)

Table 3-3. Basic equation to determine the number of theoretical separation stages Nf and the number of transfer units NTU,, of countercurrent absorbers [3.9]. Stage concept (Balance around a theoretical stage)

HTU, NTU-concept (Balance around a differential height element)

Concentration change

AY; = Y, - Y;*

dY; = ( Y ;- Y ; * ) .dNTU,,

Mass balance

6,. AY; = L,. AX;

Enthalpy balance *

L,. cp,l * A19 = G,. A Y ; . AfiAD,;

6,. dY; = L,. dx; L,. cp,, . d19 = G,. d Y i . AhAb,;

Phase equilibrium

*

Absorption heat is only transfered to the liquid phase, heating of the gas phase may be neglected. i Absorbate component, key component cp,, Specific heat of solution A , , B, Constants

3.5 Countercurrent Phase Flow Absorption, Design of Countercurrent Flow Columns

251

'S

xu X--b

Fig. 3-7. Determination of the number of theoretical stages in an operating diagram of a countercurrent absorber including equilibrium curve and balance line, depending on the solvent ratio and recirculation ratio. BL Balance line for the case of a single solvent pass through the absorber ( r = 0; v = tanx) BLM Balance line for the case of a minimum solvent ratio and a single solvent pass through the absorber ( r = 0; v = vmin= tammin) -. BLR 1 Balance line for the case of solvent recirculation ( r = r ) BLR 2 Balance line for the case of maximum solvent recirculation ( r = rmax)EC Equilibrium curve Y Gas load X Liquid load

Y -H i X ____-.___ 1 + Y p 1+x

(3-14)

(From Fig. 3-7 it is seen that recirculation of loaded solvent decreases the amount of solvent required, but due to the lower driving forces for the transfer of absorbate leads to an increased number of separation stages in the absorption column.) The solvent ratio v is an important operating variable for absorption, similar to the reflux ratio for rectification. It controls the slope of the balance line for an absorption problem and has to be larger than a minimum solvent ratio vmin. vmin is obtained by graphical means from the slope of the balance line through the points A and S (Fig. 3-7) or calculated by

are linked by where Y, and X,,,,, Eq. (3-14). With the minimum solvent ratio vmin the required degree of absorption could only be carried out with a n infinite number of separation stages. Chosen solvent ratios for practical absorber operation are usually in the range v = (1.3 . . . 1.6) vmin. The operating costs of absorption are increased with increasing solvent ratio and therefore a higher consumption of solvent. The investment costs as a consequence of a decreasing number of theoretical separation stages of the absorber with an increasing solvent ratio are decreased. Therefore an economical, favor-

252

3 Absorption

able solvent ratio voptis gained from a diagram where operating, investment and total absorber costs are plotted against v. vOptis the solvent ratio which gives a minimum for the total costs. In the concentration range Y, > Y > Y, with a constant solvent ratio v and a constant slope rn for the equilibrium curve, the absorption factor A is constant (3-16) Then the number of theoretical stages can be calculated with

In A

(3-17)

Absorption is especially economical with absorption factors A between 1.25 and 2.0. Determination of the number of theoretical stages for desorption using countercurrent solvent flow stripping, by means of a stripping gas is analogous. In this case, the fact that X, is the entry concentration and X , the exit concentration of the solvent to the stripper must be considered. According to the operating conditions, the course of the desorption equilibrium curve is different from that for the absorption equilibrium curve. The balance line is now below the equilibrium curve and so the reciprocal value of the absorption coefficient is required. The relation for the stripper is similar to Eq. 3-17

X u - Xa -

--x, Ya rn

?;( );(

A

-1

(3-19) is the ratio of the actual amount of scrubbed absorbate to the maximum possible amount when phase equilibrium between the scrubbed gas and feed solution has been reached. The scrubbing degree EA represents the yield of the absorption achieved. EA corresponds to the absorption coefficient with A < 1 and a large number of separation stages Nt. With A > 1 and a large number of separation stages N,, the scrubbing degree is approximately 1. This means that the amount of absorbate absorbed is the maximum possible. The scrubbing number

Ez =

1 Ni+ 1

where m is the slope of the desorption equilibrium curve, Nt is the number of theoretical separation stages of the stripper and X,, X,, Y,, Y, are the loadings according to Fig. 3-4. Fig. 3-8 gives an estimate of the number of theoretical stages required for absorption and stripping as a function of the absorption factor and the residual gas load. Since data in Fig. 3-8 are calculated by means of Eqs. (3-17) and (3-18), which are based on linear equilibrium curve and balance line, an appropriate concentration scale for the solvent and gas has to be used. Apart from the absorption factor A , the scrubbing degree EA and scrubbing nurnber Ez are also important. The scrubbing degree

(3-18)

Y , - rn . X,

Y, - rn X,

(3-20)

is a measure of the investment expenditure for the absorption. The smaller the difference Y, - rn. X, at the top of the ab-

3.5 Countercurrent Phase Flow Absorption, Design of Countercurrent Flow Columns

253

1.o 0.8

0.6 0.L

0.3 0.2

t

0.1 0.08 0.06

-75 0.01 'C c

0.006 0.OOL

0.003

0.002 0.001 0.0001

0.000~ O.OOO!

Fig. 3-8. Number of theoretical stages of absorbers and strippers for the case of linear equilibrium curve and balance line, and constant absorption coefficient. A Absorption parameter 1/A Stripping paramter Nl Number of theoretical separation stages

sorber, the closer is the desired approach to the equilibrium between the remaining gas and feed solution, the larger must be the number of separation stages of the absorber. Basic equations for mass and enthalpy balances, phase equilibrium relations, and the formulation of the concentration differ-

ences are listed in Table 3-3. These equations form an algebraic equation system which is the basis for stage-to-stage calculation methods for a computer [3.22-3.241. These equations are extended, if necessary, by the inclusion of stoichiometric conditions if mass or mole fractions are used for the concentration scales.

254

3 Absorption

The number of actual stages Np or trays in an absorption column is

Egm =

(3-22)

the tray efficiency is found experimentally under conditions as close as possible to the operating conditions. (To calculate E,, see Chapter 2.5.6.1). The height of the absorption column 2 for mass and heat transfer is given by

(3-21) where E,, is the mean tray efficiency for the gas phase in the concentration range of interest. (Direct graphical determination of Np may be carried out using an operating diagram as shown in Fig. 3-7. The separation stages can be drawn between the operating line and an auxiliary line, which is constructed by dividing the distance Y - Y* into the ratio 1 : Egm,y for each stage. is the efficiency at Y on the absorption concentration profile.) According to Chapter 2.5.6 and Fig. 3-9,

Z = Nt * HETS

(3-23)

where HETS is the active column height with a separation effect equivalent to one theoretical separation stage. HETS is also found by experiment under conditions as close as possible to the operating conditions (see Chapter 2.5.6.2).

b)

a)

xn-1

y , - Yn-1 Y, - Y;-1

cu

*I

i I

I

yn

I I Xn-f

Xn

X-

Fig. 3-9. Determination of the tray efficiency Ex, in an absorption column. a) Schematic b) Operating diagram BL Balance line EC Equilibrium curve X,-], X,, Absorbate load of solution found experimentally under conditions as close as possible to operating conditions Y,- I , Y, Related absorbate load in the gas phase Y:-, Equilibrium load at X,, Y Absorbate load of the gas mixture X Absorbate load of the solution

3.5 Countercurrent Phase Flow Absorption, Design of Countercurrent Flow Columns

If the HTU-NTU concept is to be applied to determine the column height Z (height of filling material, packing height) for mass and heat transfer then according the discussion in Chapters 1.7 and 1.9,

255 (3-26)

and (3-27)

dY

z=kg,Y GaeT .AQ. Ys,- Y - Y " *

= HTU,,

9

NTUog

(3-24)

or

where j3g,y, PI,, are the mass transfer coefficients for the gas and liquid phases and m is the slope of the equilibrium curve. In the range of validity of Henry's law (see Chapter 1.4.3.3) and under small loadings, (3-28)

= HTU,,

*

NTU,,

(3-25)

where

.

.

carrier gas flow rate, solvent flow rate (kmol/h) volume specific effective mass ae and heat transfer surface (corresponds nearly to the true wetted packing material surface) column cross section AQ kg,Y* k1,X overall mass transfer coefficient related to the gas and liquid phases absorbate loading of the carrier x,Y flows. (Under low key component concentrations, which is common in absorption practice, the loadings are approximately the mole fractions, y = X x = X . If this simplification can not be made, molar fraction are expressed as concentration scales in Eqs. (3-24) and (3-25).) GT, LT

According to Chapter 1.7.3.1 the mass transfer coefficients are given by

where Hi is Henry's constant and p is the absorption operating pressure. Pg,y and PI,, are related to the loading, and bgand PI to the molar fractions, the conversions are

where eg,el are the mean densities and Mg, M, are the mean molar masses of the gas and liquid phases. NTU,, and NTU,, are found by numerical or graphical evaluation of the integral expressions in Eqs. (3-24) and (3-25) (see Chapter 1.6.3). For an approximately linear equilibrium curve and balance line, the approximate result according to Eq. (1-218) applies. The heights of the transfer units HTU,,, HTU,, in Eqs. (3-24) and (3-25) are strongly governed by the total mass transfer coefficients kg,y, k , , and the specific transfer area a,. Therefore, the product kg,y.a, and hence HTU,, or kl,x. a, and HTUoI,should be found by experiment under conditions as close as possible to the operating conditions for scale up purposes (Fig. 3-10).

256

3 Absorption

Fig. 3-10. Height of a transfer unit HTCJogfor different Montz packings with regular geometry as a function of different gas and liquid loads. Representation according to Montz GmbH. Absorptions system: air, ammonia/water f Specific liquid load

a) Gas load effect 0 ,* i = 10.55 m3/(m2+h) + i = 15 m3/(m2. h) 0, A i = 10 m3/(m2. h)

Z p (m) (m) @bar) 0.29 1.36 998

Symbol Montz-Pak d O

*,

A

B1-100

+ c1-200 B1-300

0.29 1.43 992 0.22 1.0 991

b) Liquid load effect 0 ,* F = 1 r i a A F = 1.6m

257

3.5 Countercurrent Phase Flow Absorption, Design of Countercurrent Flow Columns

Calculation of the overall mass transfer coefficients is done using Eqs. (3-26), (3-27), and (3-29) after computation of the mass transfer coefficients Pg and P,. For example, to calculate the mass transfer coefficient PI of the liquid, empirical correlations are applied in the form

-

Sh, = C, Re;" . Sc; . Gap

(3-30)

where the dimensionless numbers are PI . L, Sh, = ___

Dl

Sherwood number of the liquid phase

L, characteristic length

D,diffusion coefficient in the liquid phase Re, =

l*L, ~

Dl

Reynolds number of the liquid phase

I sprinkle density, volumetric solvent flow rate of the column cross section v, kinematic viscosity of the solution VI

Sc, = - Schmidt number of the liquid Dl

phase

Gal = * " Galilei number v:

The calculation for the gas phase mass transfer coefficient, Pg using empirical correlations are of the form =

C,

. Re,' - Sci

Dg

Schmidt number of the gas phase

Dg diffusion coefficient in the gas phase Suitable correlations to calculate the mass transfer coefficient for absorption and desorption are listed in Table 3-4. Figure 3-11 shows the product of the volumetric mass transfer coefficient Pg and effective volume specific surface for mass transfer, a, as a function of gas and liquid loads for different types of random packings, found by empirical means. If a chemical reaction takes place between the absorbate A and absorbent B (chemisorption), the equations listed in Table 3-4 to calculate the mass transfer coefficient may be applied, especially when the resistance to mass transfer is mainly in the gas phase. However, if the mass transfer is controlled by the mass transfer resistance in the liquid phase, the mass transfer coefficient of the liquid phase is increased in the case of fast chemical reactions. This increase is expressed by the enhancement factor E which is defined as the ratio of the liquid phase mass transfer coefficients, P / R with and P, without chemical reaction

(3-32)

g acceleration due to gravity

Sh,

V

Sc, = -K

(3-31)

where the dimensionless numbers are

w -L , Reg = - Reynolds number of the vg gas phase

E depends on the order of the reaction, the reaction rate constant k and the diffusion coefficient D of the reactants. For a first order irreversible reaction, given, e. g., by [3.34], (3-33) In absorption practice a second order irreP versible reaction of the form A + v,B frequently appears. E is then found from the Hatta number Ha (Fig. 3-12) where -+

w superficial gas velocity vg kinematic viscosity of the gas phase

4. Correlations to calculate the mass transfer coefficient [3.68]. 4a. Liquid phase mass transfer coefficient

Packing type

2 . Ref.59. S C ~. .Gay.“ ~ 5 . Rep” . Sc?’

. Gaf.17

PI.

Characteristic length L, in

L,=4

Berl saddle

L, =

L, = dp

Raschig rings

Sh,

Re,

6

Lc=4

Gal L, =

4

Lc=d,

L, = l/ab*’ Raschig rings . Ref’3 . Sc-’.’ . (a . 4)0.4 Pall rings Berl saddle Cylinder

:I

__ 9)’12

.

( .i. . (T

d~

@I

Volume Validity range with respect to specific packing dimension surface of the packing [mm] Re, a

a

ab*)

10 5 dp 5 50 10 5 dp 5 50

6 5 dp

72

3 5 Re, 5 3 . lo 3 5 Re, 5 3 . lo

0.4 5 Re, 5 lo3

Ceramic packing

Raschig rings Saddles Spheres

a

nal diameter of packing (mm)

1

ecific wetted packing surface

L, = - Valid for all dimensionless numbers E

~

g

We,=

L, . w2. @

Weber n

(T

Relative void fraction

Fr, =

i . L,

(T,

Froude number, liquid

Critical surface tensio complete wetting [cos 0 = 1 (0Contact

Gas phase mass transfer coefficient Packing type

. ’

.

for for

2

. S CE” ~ . (a .$,)-2

CL,2 15 mm

4 < 15 mm

__ 3 1 ’ 3

w E

Raschig rings Pall rings

p,. Characteristic length Lc in

L = 4 . -&

E Lc = 4 . a

Sh,

Re,

1 Raschig =C rings a Berl saddle Spheres Miscellaneous packings

a

1

Volume Validity range with respect to specific packing dimension surface of the packing [mm] Re, a Raschig rings 10 5 Re, 5 lo4 10 5 dp, 5 50 Pall rings a6 *)

L, = -

4 = 50 10 s dp s 50

1 5 Re, 5 lo3

R

[

[

a

[

. dp > 1000 . V,

7 Raschig rings (ceramic) 30 Saddles (ceramic) 73 Spheres (ceramic)

. ent

. Sc0.33 2

,

. R e y 2 . Sc0.33 rcurrent 4

. w,

__

-

a va

, Sh,=- p g

’

[

Tower with packed tubes Pall rings Tower with packed tubes Pall rings dh

Dg

draulic diameter of packing

260

3 Absorption

a)

[s-’I 1

2

1

-

3 6 8 1012 ;.lo3 ~ r n ~ / ( r n ~ . s ) ~

2

1

3 L 6 /.lo3 [rn3/(rn2.s)l

8 1 0

2

LO

d)

t

2o

pg .a, .lo3 [s-’I10 8

6 I

I

I

I

I

I

~ I ~ I

1

2

L

6

8 10

15 20

Fig. 3-11. Gas phase volumetric mass transfer coefficient p g .a, as a function of the liquid rate. [3.47]. Representation according to BILLETand MACKOWIAK a, b) System: Ammonia/air-water, p = 1 bar, T = 288 K, d = 0.3 m a) Packing height 0.9 m, F = 1 . 2 1 p a (Hiflow ring 25 mm, PP, 45500 l/m3, F =1.621/% (Pall ring 25 mm, PP, 55 180 l/m3) b) Packing height 1.35 m, F = 1.161/Pa, 50 m m Pall ring, PP, 6425 l/m3; 50 mm Hiflow ring, PP, 6450 l/m3 c, d) System: Carbon dioxide/air-water, p = 1 bar, T = 295.9 K, F = 0 . 5 5 m , d = 0.3 m c) Packing height 0.85 m, 25 mm Hiflow ring, plastic, 46100 l/m3; 25 mm Pall rings, plastic, 55180 l/m3 d) Packing height 1.35 m, 50 mm Hiflow rings, PP, 6400 l/m3; 50 mm Pall rings, PP, 6765

i Ha

urn3

Specific liquid rate

=

reaction rate mass transfer rate (3-34)

and a parameter ERi= 1

+-

CB

v B * ‘AG

(3-35)

where cAGis the concentration of the absorbate A in the phase boundary, and cB the concentration of the reactant B in the liquid phase. The effective specific surface for mass and heat transfer, a, is mainly found experimentally, along with the mass transfer coefficient. In general a, does not correspond either to the possible theoretical geornetric specific surface a of the packing (see Tables 2-26, 2-27) or to the true wetted

3.5 Countercurrent Phase Flow Absorption, Design of Countercurrent Flow Columns

I

261

E 100

10

1

0.1

1

100

10 Ha

1000

Fig. 3-12. Enhancement factor E as a function of Hatta number Ha, as a ratio of the reaction rate and mass transfer rate of an irreversible 2”dorder reaction (parameter ERi according to Eq. (3-35). Ha Hatta number E Enhancement factor

specific surface a b . The wetted surface ab is normally smaller than the geometric surface a because of stagnant flow volumes and zones as a result of poor wetting, especially in random packing. Again a, is smaller than ab, due to the formation of coherent liquid zones in packing elements or liquid gussets. (The additional defined hydraulic surface is the sum of the nonwetted surface and the surface for mass transfer of the packing.) The active volume specific phase interfacial area for mass and heat transfer, a, depends predominantly on the sprinkle density v//Ae, the properties of the liquid phase and the wettability of the packing surface. A small sprinkle density (e.g., in vacuum rectification operation), large surface tension CJ and high viscosity ql leads to a small a,. The poor wettability of plastic packing material to aqueous liquids or other polar liquids is improved by hydrophilizing the surface of the packing material [3.42]. Only with column operation above the lower loading point (see

Fig. 2-75) the gas load is also important, which increases a,. An estimate of the surface area of the packing, a, for mass transfer by physical absorption according to PURANIKand VOGELPOHL[3.43] for the liquid phase is

- I

(3-36)

including the Reynolds number

i

Re, = ____ vI. a

(3-37)

and the Weber number (3-38) The “critical” surface tension uCfor wetting is for packing material steel 71, pottery 61, PVC 40 and glass 73 N/m. More

262

3 Absorption

calculation methods for a,, are given in [3.44-3.461. For further design and optimization of absorption units with countercurrent flow and solvent regeneration, see [3.4, 3.253.301. The calculations for adiabatic operated absorbers, in which large heat effects have to be considered, are described in [3.31]. Process control of absorption columns can be found in [3.32]. The dynamic behavior of absorption tray columns is discussed in [3.33].

3.6 Types of Absorber In principle all types of gadliquid contact apparatus are suitable for use as absorber. For selection it has to be ensured that 0

0

0

0

The available driving force for the absorption is used to its maximum The minimum input of conveying or mixing energies (energy dissipation) obtains a maximum phase surface area with good surface renewal With high phase turbulence, mass transfer coefficients are large Special absorption problems such as unusual loading, treatment of corrosive substances, possibility of incrustation, strong heat effects etc., are considered

flow scrubber; scrubber with rotating internals such as rotation, cross haze, and disc scrubber) 0 Gas comes into contact with the liquid, which has the form of a film on a fixed or moving base, gas and liquid phase remain continuous (column absorber with random or regular packing, falling film or surface absorbers) Some important absorber types are discussed in Table 3-5 (see pages 264-273) with references made to particular literature. Figure 3-28 compares selected absorbers with respect to the volumetric specific phase surface a as a function of the specific conveying or dispersion power input, which is

lobT

L

I

I

1

1- --1

rn3

Gas and liquid come into contact in three different ways : 0

0

Gas is dispersed into the wash liquid, the liquid phase remains continuous (column absorber with transfer trays, bubble columns, vessel absorber, dispersing agitator) Wash liquid is sprayed into the gas, the gas phase remains continuous (free board scrubber without internals such as venturi srubber, injector srubber, spray srubber, annular-flow srubber, radial-

O2

w[5] -

Fig. 3-28. Specific contact surface of different absorbers as function of specific power input. Representation according to MERRSMANN, HOFER,STICHLMAIR [3.9]. a Volumetric or specific interfacial area W Specific power input

3.7 Regeneration of the Solvent, Desorption

where A is the phase interfacial area between gas and liquid, V, is the absorber volume, W is the induced conveying or dispersion power; C, q and r are factors depending on the absorber design; 1 - E is the relative gas content and E is the relative liquid content [3.34-3.361. Packed columns are mainly employed in absorption processes. They are distinguished by having a low pressure drop (ca. 2-5 mbar/m), allowing a variable gas load and - similarly to tray columns - enable high numbers of separation stages. Spray scrubbers with one or at the most a few separation stages, are employed as chemical scrubbers if small residence time is sufficient. The pressure drop is very low at ca. 0.1 mbar/m. Venturi scrubbers with a liquid atomizer have a high gas pressure drop but provide very intensive phase contact. They are particularly suitable for scrubbers with short residence times in absorption processes and, in combination, for fine dust separation. Bubble columns, dispersion agitators and vessel absorbers, have a relatively large liquid volume, but small gas loading and a high pressure drop. Since the residence time of the liquid is large, slow running absorption processes suit these apparatus [3.9].

3.7 Regeneration of the Solvent, Desorption Absorption is favored at raised pressure and low temperature. Therefore, the reverse process, desorption, favors low pressure and high temperature. With desorption the absorbed component is removed from the absorbent; the solvent is degassed and regenerated before reuse. On the whole, desorption is carried out in four ways, which can either be applied individually or in combination:

263

Simple expulsion of the absorbed gas component of the absorbent by reducing the pressure. This desorption option is advantageous if the absorption was carried out under a higher operating pressure 0 Expulsion of the absorbed gas component under higher temperature, usually linked to rectification and combined with pressure release steps. The boiling point of the solvent limits the operating temperature 0 Expulsion of the absorbed gas component by an inert gas or steam flow (stripping) which is mainly carried out in packed or tray columns. Inert gas flows countercurrently to the loaded solvent. The absorbed gas component travels from the liquid phase into the gas phase, in which its partial pressure is kept low by the continuously fed inert gas 0 Formation of a chemical compound between absorbed component and a regeneration auxiliary substance (chemical regeneration). The absorbed component reacts with the auxiliary substance which is insoluble with the absorbent or solvent, and precipitates. Downstream to the regeneration unit is a thickener, and the solid is separated from the absorbent in the subsequent filter (e. g., the double alkali process to separate sulfur dioxide from flue gas. Sulfur dioxide forms a sulfite with an alkali solution used as an absorbent. When calcium hydroxide is added calcium sulfite is formed, which is insoluble in water and therefore precipitates. This is then separated from the alkali solution). Some regeneration processes for loaded solvent are shown in Fig. 3-29. 0

Another example for the separation of an absorbed component from the washing liquid (regeneration) is given in Fig. 3-30. In the presented Wellmann Lord process, sulfur dioxide is separated from flue gas using

. Absorber types (type, assembly, and operation principle, references). le column [3.48-3.531, vessel absorber

le columns without liquid circulation ') stage

ple bble lumn

Bubble column cascade

b) tered te

Jltistage -

(Sieve t r a y ) (Single hole t r a y ) Bubble column with cascade tray

Columns with gas distributor (porous sintered plate, plate, sieve tray, spigots, nozzles), gas flows cocurrently currently through a continous liquid phase in form of gas jets; homogenous (quasi-laminar) bubble flow, is ch by a large interfacial area and little backmixing with equ at a superficial gas velocity wg < 0.05 m/s. Bubble coal form larger bubble, bubble splits, heterogenous (turbule flow may be observed at wg > 0.05 m/s. wg,,, = 1 m/s; ing range ca. 0.01 -0.2 m3/(m2 . s); ratio height/diamet diameter 0.1 - 10 m; in large dimension and inexpensive t high gas flow rates possible at large liquid residenc movable internals. Fig. 3-13. L liquid phase, G gas phase.

a)

Perforated plate

Gas ____

-200 prn

Sieve tray

-Gas

+

One-hole tray

Liquid-

depending on

0.5-5 mrn

tg/f,

Fig. 3-14. Devices used to generate bubbles'). db Hole diameter Volumetric flow rate of gas and liquid phase

c,

0’lj--j;~

columns with liquid circulation ’) l tion G

0

Forced

0

LQ

othir mp)

slip stream-

Jet-

L Jet tubereactor

L

b)

mixing

d)

C)

Ring Furminozzle gating

1

Ejector nozzle

0

G

Recycle reactor [3.48, 3.543 with natural circulation t central tube (mammoth recycle reactor), with cocu stream of liquid and gas phase (slip stream recycle re internal tube (jet tube recycle reactor) (Fig. 3-15 a-c).

Jet tube reactor [3.55, 3.561 with momentum exch (“ejector nozzle”) above the ring nozzle to inte distribution and to keep the gas in the range of the (Fig. 3-15 d). Jet reactor [3.48, 3.561 with a gas distribution by a t nozzle and liquid circulation preferential in the low section (Fig. 3-15e).

Jetreactor

Recycle reactor ___ )

zle

ber

Fig. 3-15

el

Venturi

3-5d-d-

tube


(100

(20

10-30

10-30

0m/s

Fig. 3-16. Nozzle systems to generate bubbles’). w ~maximum , ~ liquid ~ exit ~ velocity

5-20 c3

(continu

. (continued)

l absorber (bubbling column)

CG

-CA

FS

FS

q j f 000

v

OD0

v

000

v

000

v

GO

-

v

LS

CA

LS

In bubbling columns [3.1] during the absorption proces heat transfer takes place at the interface of bubbles ri wash liquid, which is at rest or slowly moving. Bubblin are suitable when poorly liquid-soluble gases have to b (Fig. 3-17).

Fig. 3-17 FS Fresh solvent Loaded solvent Crude gas Purified gas Gas distributor Cooling agent

LS CG PG GD CA

g plate reactor, mixer, gadliquid reactor system ER[3.57-3.591.

iquid

I-

In a lifting plate reactor2) fed gas is dist lifting elements (a perforated disc packag on a central lifting shaft, disc distance c hole size 12 mm, hole spacing ca. 27 amplitude 100 mm, lift frequency ca. 1 mixing reactor2) gas is dispersed by bl made of 6 or 12 blades (Fig. 3-18b). Gadliquid reactors3) made of single stag ing dispersion zone for phase mixing by mixer and phase separation zone, suitable rent and countercurrent operation (gas ho 20%, stage efficiency 48-97%, measured carbon dioxide-water system) (Fig. 3-18 a). type is suitable to bring a small gas flow with a large liquid flow, where the transfer resistance is in the liquid phase fore, a large liquid residence time is req

Fig. 3-18 1 Reaction stage 2 Demixing (separation) stage 3 Agitator shaft with perforated disc mi

(continued)

pe absorber and surface absorber [3.1, 0.61

(Surface absorber) (Fig. 3-19). Absorption of easily solub an absorbent moving slowly in thick layers in vessels (touri or plates. Simple design. Suitable for corrosive substances absorption enthalpy.

absorber

(Film-type tube apparatus): Multitube apparatus with a bent falling-film on the tube insides. Cocurrent and count flow of gas phase and falling-film (both phases are cohere ple design with the possibility of heat removal by a cooli outside the tube shell. Low gas pressure drop, small interfa small liquid mass transfer coefficient.

pe absorber

Fig. 3-19 ent

f Crude gas

column (falling-fiIm column) column (Figs. 2-70a and 2-71)

Column with random packing (see Table 2-27) or tower dom packed tubes (heat exchange possible to a cooling the shell side). Specific surface: 50-300 m2/m3; diamet 2.5 m; vapor load factor: 1-2.5 Pa”2; trickle density: 2(m2 . h); specific pressure drop: 20- 120 mm water gauge; liquid phase are both continous. Phase contact may also take place at movable or fluidiz material (usually light-weight spheres). The contact app then a fluidized-bed contactor or floating bed contact Specific surface: 150-250 m2/m3; superficial gas 2.8-4.2 m/s; trickle density: 40-150 m3/(m2. h).

cked column (Fig. 2-70b)

n

bber

rified gas

Spray scrubber

Column with regular structured packing (Table 2-28). Specifi face: 100-500 m2/m3; gas load factor: 2-5 Pa”’; trickle de 0.6-200 m3/(m2. h); specific pressure drop: 5-40 mm gauge; gas and liquid phase are usually continous. Columns with trays with - or seldom - without controlled l flow and, therefore, a stagewise or continous phase c (Table 2-18-2-20.). In the two-phase area gas is dispersed in bles. Column diameter: up to 10 m; gas load factor: 0. Pall2; tray pressure drop: 20-80 mm water gauge.

In spray scrubbers absorbent is sprayed (dispersed) into th phase. Dispersion equipment are usually nozzles mounted tically on top of each other or horizontally in a row. Liqui gas phases are in cocurrent or countercurrent flow (Fig. Spray scrubber examples are annular gap scrubbers and radial scrubber.

-

I-

Absorbent

Fig. 3-20

absorbent

(continued n

(continued)

r gap scrubber [3.60]

In the throat of a ring gap liquid is dispersed in droplets flux. By means of a displacer the ring gap throat are adjusted to the changeable gas flow to ensure a con velocity (Fig. 3-21).

Fig. 3-21 DP Displacer DM Demister AN 1, AN 2 Absorbent flow Stage 1 and 2

-flow scrubber [3.60]

CG Crude gas PG Purified gas LA 1, LA 2 Loaded absorb from stage 1 and 2

Radial gas phase flow from the center outward in the formed by two discs. To mix intensively wash liquid is spr trally by means of a nozzle central into the gas flow. The thr ing in the center of the wash zone may be adjusted to the by means of the check ring. To improve droplet separatio flow is induced by baffles mounted on the disc rim (Fig.

Fig. NO BA AD WL

3-22 Nozzle Baffle Adjusting equipment Wash liquid

RG PG LA CR

Raw gas Purified gas Loaded absorbent Check ring

ber, Venturi scrubber [3.61-3.631

b)

Pressurized wash liquid is sprayed in a jet scrubber, and di into the slow flowing gas phase. The driving force is induce washing liquid; the gas phase is entrained by the liquid jet droplet swarm (Fig. 3-23 a).

WL

In a venturi scrubber the driving force is the gas phase. L sprayed into the gas stream via a nozzle (Fig. 3-23 b). Ab with liquid dispersion are suitable for the absorption of soluble gases with short wash liquid residence times. This a type is the most effective but requires high conveying ener therefore, high operating costs.

-\

I

1

TP

bsorber (absorption machines)

Fig. 3-23 a) Jet scrubber b) Venturi scrubber RG Feed gas, raw gas WL Wash liquid TP Two phase mixture

In a rotary absorption apparatus scrubbing liquid is disper the gas stream by rotating dispersing elements (funnels, dis etc.). Rotary absorber types are costly to construct and ener sumption is high. However, rotary absorbers are suitable f gas throughput rates at a small pressure drop. Rotary absor relatively insensitive to incrustation by dust-laden gases.

(continued

(continued)

crubber [3.60, 0.61

-MS

LA

Column with rotor. Due to the centrifugal force solvent liquid is drawn into funnels mounted in baskets on (Fig. 3-24). Wash liquid is dispersed stagewise into the u gas phase. Diameter: up to 4m; height: up to 13 m ; rotor re 70-80 min-I; number of stages: 3-10; gas throughpu 25000 m3/h; power consumption: up to 80 kW; spray ca each stage: up to 550 m3/h; droplet diameter: 1-3 mm.

Fig. 3-24 RG Feed gas, raw gas PG Purified gas SO Solvent LA Loaded absorbent RO Rotor with funnel shaped baskets DI Dish MS Mist separator

RG

r cross-fog scrubber [3.60, 0.61

Rotary absorber with two horizontal shafts. Counter channelled discs are mounted on the shafts and disper from the bottom into the gas phase (Fig. 3-25). Housing diameter: 0.5-3 m; length: 2-3 m; number o tions: 500-600 min-'; disc diameter: up to 0.5 m; numb pairs: 11-15; spray capacity per disc: 8-10 m3/h; gas thr 1000- 20000 m3/h; power consumption: 4-7 kW.

Fig. 3-25

llUllLullLal s11a11.

uas

allu

llqulu

pllabt.

arc: CrllLlally cllal

conical sieve and mixed in parallel flow in the space be metal strips. After the mixture is directed to the outside, t liquid phases are separated in the shell (Fig. 3-26). Gas put: 10000 m3/h; power consumption: ca. 15 kW ( example).

Fig. 3-26

PG

-spray scrubber [3.60, 3.651 RO

Rotary absorber with a vertical rotating spray cone. The into liquid and transports liquid upward by means of a pa liquid is dispersed in the gas phase as it exits the c (Fig. 3-27) via the cone drillings. Fig. 3-27 Rotor with spray cone Mist separator Feed gas, raw gas Purified gas Solvent Loaded absorbent

RO MS RG PG SO LA

entation according to GERSTENBERG [3.48]. entation according to MERSMA", VOIT,ZEPPENFELD [3.57]. entation according to Montz GmbH.

274 01

3 Absorption Clean gas

D

Q

c)

Clean gas

-4 Absorbate

bl

Clean gas

d)

Clean gas

P

h

Y

Absorbate

Auxiliary substance

nr-+

Fig. 3-29. Desorption processes variations. Representation according to STICHLMAIR [3.67]. a) Multistage pressure release of the loaded solvent b) Regeneration by rectification c) Steam stripping d) Regeneration by precipitation

3.7 Regeneration of the Solvent, Desorption

b'

I I I

I

I i I I I I I I I I I I

U

I

-----

t

I

L

0)

c

6In P 0

+

-0

+ 0

r

t-

A

' 0 I I

275

ater

-r6

s0,-

rich

"$

10

ter

~

I -

M-

r-

Soda(NaOH1

11 Final product

. Flow chart (a) and simplified flow sheet (b) for flue gas desulfurization by the Wellmann Lord process. ntation according to Davy McKee, Zimmer AG, Frankfurt/Main [3.66]. gas blower; (2) Heat exchanger; (3) Prewasher; (4)Absorber; (5) Filter; (6) Vessel; (7) Evaporator; (8) Condenser; (9 arator; (11) Dissolving vessel; (12) Sulfate crystallizer; (13) Centrifuge; (14) Dryer (alternative).

3.7 Regeneration of the Solvent, Desorption circulated, concentrated sodium sulfite solution, the absorbed component forming sodium hydrogen sulfite Na2SO3+SO,

+H,O

absorption .=

' 2NaHS0,

regeneration

During the subsequent thermal regeneration of the circulated solution in an evaporative crystallizer, water is vaporized, and due to t h e shift in the equilibrium at high temperature, sulfur dioxide is evolved. After condensation of the water in a condenser, the desorbed sulfur dioxide is a highly concentrated gas. Redissolving the sulfite crystals from the crystallizer with condensed vapor produces fresh wash liquid which is the reflux to the absorber. The basic relationship for the design of a desorber is analogous t o that of the design of an absorber. The reverse mass transfer direction and the different phase equilibrium situation must be noted.

References [3.1] THORMANN, K. : Absorption. Springer, Berlin 1959. [3.2]RAMM,W. M. : Absorptionsprozesse in der chemischen Technik. Verlag Technik, Berlin 1952. [3.3] MORRIS,G. A., and JACKSON, J.: Absorption Towers. Butterworths Scientific Publications, London 1953. [3.4]SHERWOOD, T. K., and PIGFORD, R. L.: Absorption and Extraction, McGrawHill Book Comp., New York 1952. NORMAN, W. S.: Absorption, Distillation and Cooling Powers. Longman, London 1962. KOHL,A. L., and RIESENFELD, F. C.: Gas Purification. McGraw-Hill, New York 1960.

277

[3.7] DIETER,K., and HUBNER,W.: ,,Gasabsorption. " Fortschritte der Verfahrenstechnik, Vol. 8 (1966/67).Verlag Chemie, Weinheim 1969. [3.8]MERSMANN, A.: Staub Reinhalt. Luft 36 (1976)8, 331. [3.9] MERSMANN, A., HOFER,H., STICHLMAIR, J.: Chem. Zng. Tech. 51 (1979) 3, 157- 166. [3.10] STICHLMAIR, J. : ,,Absorption und Rektifikation. '' Fortschritte der Verfahrenstechnik, Vol. 20 195-220, VDI-Verlag, Dusseldorf 1982. [3.1I] BILLET,R., and MACKOWIAK, J. : ,,Rectification and Absorption. " Fortschritte der VerfGhrenstechnik, Vol. 18 239-269. VDI-Verlag, Dusseldorf 1980. [3.12] MOLZAHN, M., and WOLF,D.: Chem. Ing. Tech. 53 (1981) 10, 768-780. [3.13] EMBEGER, J., KERN, H., LEMPP, M., SATTLER, K., and STAHL,R.: Umweltschutz, Entsorgungstechnik. VogelVerlag, Wurzburg 1982. [3.14] FISCHER,H. J.: Chem. Tech. Heidelberg 10 (1981)4, 297-300. [3.15] cav 1984, 3, 22, 102. [3.16]International Critical Tables. McGrawHill Book Comp., from 1933. [3.17] HERPERS, E. T.,and DUERKOP,A. TECH. MITT. 75 (1982)2/3, 152-158. [3.18]K~~~,H.:Chem.Znd.37(1985)5,349-353. [3.19]FORCK,B., and LANGE,G.: Systemanalyse Entschwefelungsverfahren. VGB Technische Vereinigung der GrolJkraftwerksbetreiber, Essen 1975. [3.20]FISCHER,H. : ,,Schwefel." Ullmanns Encyklopadie der technischen Chemie, Vol. 21. Verlag Chemie, Weinheim 1982. SANDER, U.,ROTHE,U., and KOLA,R.: ,,Schwefeldioxid, Schwefelsaure." Ullmanns Encyclopadie der technischen Chemie, Vol. 21. Verlag Chemie, Weinheim 1982. [3.21] KAMINSKI,W.: Chem. Ing. Tech. 55 (1983)9, 667-683. [3.22] BURNINGHAM, D. W., and OTTO, F. D.: Hydrocarbon Process. 46 (1967) 10, 163- 170. .~~ .. 13.231 GELBE,H., and NOMINE,H . : Verfahrenstechnik 5 (1971) 10,429-435. ~

278

3 Absorption

[3.24] FUTTERER, E., LANG,G., and NEUMANN, K. K.: DECHEMA-Monogr. No. 14101431 73 (1974). [3.25] NAGEL,S., and R E L E N B.: , Chem. Ing. Tech. 44 (1972), 6, 416-420. [3.26] HAAS,J. R., GOMEZ,A., and HOLLAND, C. D.: Sep. Sci. Technol. 16 (1981) 1, 1-24. [3.27] CHOWDHURY, M., ISHIKAWA, T., and HIRATA,M.: J. Chem. Eng. Jpn 13 (1980) 8, 548-552. [3.28] OWENS,W. R., and MADDOX,R. N.: Znd. Eng. Chem. 60 (1968) 12, 14-28. D. W., and OTTO,F. D.: [3.29] BURNINGHAM, Hydrocarbon Process. 46 (1967) 10, 163- 170. [3.30] UMEDA,T.: Ind. Eng. Chem. Process. Des. Dev. 8 (1969) 3, 308-317. M. K. : Can. [3.31] RAAL,J. D., and KHURANA, J. Chem. Eng. 51 (1973) 4, 162-167. A.: [3.32] SCHIPPERS,H., and MERSMANN, Chem. Zng. Tech. 46 (1974) 5, 201. J.: Chem. Zng. Tech. 44 [3.33] STICHLMAIR, (1972) 6, 411-416. B. and KURTEN,H.: [3.34] NAGEL,O., HEGNER, Chem. Zng. Tech. 50 (1978) 12, 934-944. [3.35] NAGEL,O., KURTEN,H., and SINN, R.: Chem. Ing. Tech. 42 (1970) 474-479. [3.36] NAGEL,O., KURTEN,H., and SINN,R . : Chem. Zng. Tech. 44 (1972) 367-373. J.: Grundlagen der Dimen[3.37] STICHLMAIR, sionierung des GadFliissigkeit-Kontaktapparates Bodenkolonne, Verlag Chemie, Weinheim 1978. [3.38] REICHELT,W.: StrOmung und Stoffaustausch in Fiillkorperapparaten bei Gegenstrom einer jliissigen und einer gasformigen Phase. Verlag Chernie, Weinheirn 1974. W. : Chem. Tech. (Heidelberg) [3.39] REICHELT, 5 (1976) 6, 213-219. [3.40] MCCARTHY, J. E.: Chem. Eng. Prog. 76 (1980) 5 , 58-62. [3.41] BILLET,R., and MACKOWIAK, J.: Chem. Tech. (Heidelberg) 13 (1984) 12, 37-46; 14 (1985) 4, 91 and 14 (1985) 5, 195-206. [3.42] Company report: ,,Hydrophilierte Kunststoff-Fiillkorper," Raschig GmbH. [3.43] PURANIK,S. S., and VOGELPOHL, A.: Chem. Eng. Sci. 29 (1974) 501-507. [3.44] KOLEV,V. : Ve'erfahrenstechnik3 (1969) 6, 241-243.

[3.45] ZECH,J.: Dissertation, TU Miinchen 1978. A. : Thermische Wrfahrens[3.46] MERSMANN, technik. Springer-Verlag Berlin, Heidelberg 1980. J.: Chem. [3.47] BILLET,R., and MACKOWIAK, Tech. (Heidelberg) 13 (1984) 12 and 14 (1985) 4, 5 . [3.48] GERSTENBERG, H.: Chem. Zng. Tech. 51 (1979) 3, 208-216. [3.49] BLENKE,H., and HIRNER,W.: VDI Ber. 218 (1974). [3.50] HONG, W.-H., and BRAUER,H.: VDI Forschungsh. 624 (1984). [3.51] MERSMANN, A.: Chem. Ing. Tech. 49 (1977) 9, 679-691. [3.52] DECKWER,W. U., and SCHUMPE,A.: Chem. Ing. Tech. 57 (1985) 9, 754-767. [3.53] ZEHNER,P.: Chem. Zng. Tech. 54 (1982) 3, 248-251. H.:K VerfahrensE, [3.54] HIRNER,W., ~ ~ ~ B L E N technik 11(1977) 5 , 297-303. [3.55] NAGEL,O., KURTEN,H., and SINN,R.: Chem. Ing. Tech. 42 (1970) 7, 474-479 und 14, 921-926. [3.56] ZEHNER,P., and BITTINS,K.: Fortschr. Verfahrenstech. 23 (1985) Sec. D, 373393. [3.57] MERSMANN, A., VOIT, H., and ZEPPENFELD, R . : Chern. Ing. Tech. 58 (1986) 2, 87-96. [3.58] BRAUER,H.: Chem. Ing. Tech. 58 (1986) 2, 97-107. [3.591 Company report: ,,Montz-Gas/FliissigkeitsReaktor System Brauer," Montz GmbH. [3.60] MENIG, H.: Luftreinhaltung dutrh Adsorption, Absorption und Oxidation. Deutscher Fachschriften-Verlag Braun & Co. KG, Wiesbaden 1977. R., KURTEN,H., and NAGEL, [3.61] HOFFMANN, 0.: Chem.Ing. Tech. 45 (1973) 13,881- 887. [3.62] UCHIDA,S., and WEN, C. Y.: Znd. Eng. Chem. Process Des. Dev. 12 (1973) 4, 437-443. [3.63] RIPPERGER,S., and GERMERDONK, R. Chem. Ing. Tech. 56 (1984) 6, 466-468. [3.64] REICHELT, W. : Chem. Ing. Tech. 45 (1973) 1, 25-29. [3.65] VOGT, H. CZ Chem. Tech. 2 (1973) 9, 373- 374. U. : Chem. Tech. (Heidelberg) [3.66] NEUMANN, 12 (1983) 7, 21-24.

References [3.67] STICHLMAIR, J. : Staub Reinhalt. Luff 36 [3.75] (1976) 8, 337. W.: CZ Chem. Tech. 2 (1973) [3.68] REICHELT, [3.76] 2, 49-55. R. : Dissertation, TH [3.69] SEMMELBAUER, Darmstadt 1966. D. W., and HOFTYZER, [3.77] [3.70] v. KREVELEN, P. I.: Chem. Eng. Prog. 44 (1948) 529. E X. A.: [3.71] NORMAN,W. S., and SAMMAK, Trans. Inst. Chem. Eng. 41 (1963) 117. [3.72] ONDA,K., TAKENCHI, H., and O K U M ~ , [3.78] Y.: Chem. Eng. Jpn. 1 (1968) 56. A.: Chem. [3.73] ZECH,J. B., and MERSMANN, Zng. Tech. 50 (1978) 7, 549. [3.74] BRAUER,H. : Stoffaustausch einschfiejlich chemischer Reaktionen. SauerlanderVerlag, Aarau 1971.

279

BILLET,R., MACKOWIAK, J., and SUDER, S.: Chem. Zng. Tech. 50 (1978) 7, 550-551. SCHLUNDER, E.-U., and ~ U R N E F.: R , Destilfation, Absorption, Extraktion. Georg Thieme-Verlag, Stuttgart 1986. SCHLAUER,J., KRIEBEL,M.: “Absorption.” Ullmann’s Encyclopedia of Industrial Chemistry. Vol. B3. VCH Verlagsgesellschaft, Weinheim 1988. COULSON,J. M., RICHARDSON,J. F., BACKHURST, J. R., and HARKER,J. H.: “Absorption.” Chemical Engineering. Vol. 2. Pergamon Press, Elmsford, Oxford 1991.

Thermal Separation Processes: Principles and Design Klaus Sattler, Hans Jacob Feindt copyright 0VCH

Verlagsgesellschaft m h H , 1995

4 Adsorption

4.1 Principles of Adsorption and Desorption, Processes and Examples 4.1.1 Concept Adsorption [4.1-4.101 is the addition and bonding of certain components of a gas or liquid mixture to the surface of an active porous solid. At the interface: adsorbent

+ adsorbate

adsorption

yL desorption

adsorbed component/adsorbent The key component to be adsorbed in a free moving state is called the adsorbate, and in the bonded state the adsorbed component. The solid acts as an auxiliary substance and is described as the adsorbent. Adsorbed molecules are 6onded by physical adsorption (physisorption), with electrostatic attractive forces or Van-der Waals forces. Due to the small range of these forces, the adsorbent and adsorbate have a loose bond, which is easily released. With chemical adsorption (chemisorption), the bond is caused by valence forces. This chemical bond requires a greater effort of separation to regenerate the adsorbent, if indeed such separation is possible. Adsorption is a thermal separation process. It is mainly applied to separate components of low concentration from a gas mixture. The adsorbed gas component may be the desired product or an unwanted pollu-

tant, which has to separated from the gas mixture. In the latter case the adsorption acts as a “dry” cleaning process in environmental technology. Environmental pollutants are removed from emissions by adsorption. Adsorption is also used in the liquid phase to clean the liquid. The adsorbent is introduced into the agitated liquid to be purified, substances of the liquid are bonded to the surface of the adsorbent. The loaded adsorbent is separated afterwards from the liquid by filtration or centrifugation. Gas-phase adsorption is favored under raised pressure and reduced temperature. The adsorbability increases with an increasing boiling temperature of the adsorbate. Some adsorbents, for example, molecular sieves, favorably adsorb polar substances compared to polarizable or unpolar substances. The adsorbability decreases with increasing saturation for hydrocarbons with similar carbon number. Aromatic substances are better adsorbed than olefins, and these again better than paraffins. The movability and hence the energy of the adsorbate molecules are reduced when the adsorbed component (adsorbate) is bonded to the adsorbent. Adsorption is therefore an exothermic process. With physical bonding the heat of adsorption is < 40 kJ/mol of adsorbed substance. With chemical bonding, the heat of adsorption is > 80 kJ/mol of adsorbed substance and decreases with increasing extent of cover of the adsorbent surface. Accordingly, the exothermic bond energy is required to release the adsorbed substance from the surface of the adsorbent. Desorption is an endothermic process.

282

4 Adsorption

The reverse process to adsorption, desorption, is favored by a pressure reduction and an increase in temperature. With desorption, the loaded component or adsorbed component, is removed from the surface of the adsorbent. After regeneration of the adsorbent and treatment by drying and cooling, the adsorbent is ready for reuse.

4.1.2 Processes and Examples In practice, gas-phase adsorption is mainly carried out in a vessel adsorber with fixedbed packing mounted on a support grid, like a sieve. Since the treated gas phase flows continuously, a fixed-bed adsorption unit consists of at least two vessels. One vessel is “loaded”: from the flowing gas the adsorbent adsorbs adsorbate until the adsorbent becomes saturated, at which point the adsorbate is no longer adsorbed and is detected at the exit of the adsorber (adsorption phase). During the adsorption phase of the first vessel, the second vessel is considerably regenerated (desorption phase, regeneration phase). This is done by a temperature increase (temperature change), decrease of pressure (pressure change) or by

rinsing with an auxiliary desorption substance (displacement desorption). A combination of these variations also exists. The regenerated adsorbent is left to dry and is then returned to the adsorption operating conditions (cooling, pressure build up). Figure 4-1 shows schematically pressure swing adsorption (PSA) applied as the BF process to generate nitrogen from air, including notes and explanatory diagrams. Durink the adsorption phase, oxygen is adsorbed from compressed cooled air on carbon molecular sieve packing in a vessel switched to “load”, the remaining nitrogen leaves the vessel as product gas. From storage, nitrogen is then supplied to users. The time-controlled adsorption phase is achieved when the oxygen reaches a certain level in the product gas. In the subsequent regeneration phase, desorption starts by pressure reduction (pressure compensation with the previous regenerated vessel) and expulsion of oxygen. After the desorption phase, the pressure in the vessel is increased before starting the next adsorption phase. Figure 4-2 shows the participation of solvent recovery by adsorption. The LURGI-SUPERSORBON process shown in Fig. 4-3 uses steam as a desorption fluid for the recovery of water insoluble solvents. During the adsorption phase, solvent-rich

Fig. 4-1. BF-N, process (a) and pressure swing diagram (b) to produce nitrogen from air by pressure b swing adsorption. Representation according to data of Bergbau-Forschung GmbH and INCA mbH, Essen [4.12]. Adsorption cycle: Adsorption of oxygen on a carbon molecular sieve CMSN2 (loading period 60 s) Desorption cycle: Pressure compensation and withdrawal of remainder by means of a vacuum pump (desorption period 60 s)

Adsorption pressure/desorption pressure 1M0.1; 3.5/0.1; 8/1 bar Variations: Plant size 1-2000 m3,/h, variation 8/1 bar Output pressure 5-7 bar N, Purity 95 - 99.9 “70 H,O Dew point < -50°C CO, Concentration < 5 PPm Energy consumption of nitrogen production 0.2-0.7 KW/m3

--

I

Nitrogen product gas I

v c

_ _ _ _Control _ _ _ _ _ _ _ai _ I

Wastewater Pressure

Solubility in water Recovery process

Adsorption on active carbon

Adsorption on active carbon

I

I

Partially soluble

Insoluble

Treatment (water separation)

,

Separation (Phase separation]

I Enrichment condensation

I Enrichment condensation and distillation

Complet Adsor active

I Chemical drying

I

I

Decomposition

Examples

I

J r I Distillation I

Acetic ether, MEK. MIBK, MEK/toluene, THF/toluene. alcohoVtoluene

Toluene, xylene. benzene. hexane. chlorohydrocarbon

Ethanol, is acetone. T

Fig. 4-2. Solvent recovery by adsorption *. Representation according to KRILLand WIRTH[4.13].

*

Sometimes water washing is advantageous (e.g. at DMF)

4.1 Principles of Adsorption and Desorption, Processes and Examples

air flows through the active carbon fixed bed, switched to “load”, in which solvent vapor is mainly adsorbed. Clean gas leaves the vessel. If a breakthrough of solvent occurs, after a certain time and at a certain solvent exit concentration, the controller is switched to a regenerated adsorber. In the regeneration phase, the loaded active carbon desorbs by using steam in the opposite

285

direction to the loading direction. The solvent is expelled as vapor and condensed. The water and solvent are separated in a gravity separator. The active carbon, treated with ,steam has to be dried and cooled with the carrier gas containing the solvent at the start of the loading phase. If the adsorption process is operated continuously with respect to the adsorbent, Adsorber

Adsorber

Active carbon

b Solvent-rich waste air

I

I

t

I

T

Condenser

Water

Fig. 4-3. Simplified flow diagram of the Lurgi-SUPERSORBON process to separate solvent vapor from waste air. Representation according to data of Lurgi GmbH, FrankfurUMain [4.14].

Process data: Adsorbent:

Cylindrically shaped active carbon, particle diameter 4 mm, specific surface 1240- 1250 m3/g 1000-400000 m3/h Gas throughput: up to 99% Adsorption yield: Steam (2-4 bar, overpressure): 2-4 kg/kg +- 1-1.5 kg/kg for recPower consumption: tification (in case of water soluble solvent) Electric power consumption : 0.15-0.45 kWh/kg 30-60 kg/kg (heating by 30°C) Cooling water: 0.5 - 1 kg/t Active carbon losses:

286

4 Adsorption

the absorbent must pass through the adsorption, desorption, drying and cooling zones in a continuous cycle. For example, a continuous adsorption process to recover solvent, the Lurgi KONTISORBONprocess, is

shown in Fig. 4-4. The solvent-containing exhaust air is precleaned, if necessary, by mechanical means and enters through a control system at the bottom of the adsorption zone the unit. Together with the active

Secondary adsorber

-Two-stage

Ey!?-

Cooler -

Solvent Pneumatic conveyer

Air

-

4.1 Principles of Adsorption and Desorption, Processes and Examples

carbon, stored on perforated metal sheets, it forms a multistage fluidized bed where the adsorption of the solvent takes place. Cleaned used air leaves the unit at the top. The loaded active carbon falls from the lower gas distributor into a sluice tube which separates the air rich adsorption zone from the inert desorber. After the secondary adsorber, the regeneration zone, the active carbon passes through three tube bundle heat exchangers mounted on top of each other. The downflowing active carbon is heated in the two stage heaters. Steam or thermal oil is used as heating medium. Nitrogen flows countercurrent to the carbon and strips the desorbed solvent, which is then released when the carbon is heated. Solvent is condensed and removed as liquid in an external condenser. Nitrogen saturated with solvent passes from the condenser into the secondary adsorber, where the active carbon moves down and adsorbs the solvent. Nitrogen is recycled to the sec-

ond heating stage with a blower. The regenerated active carbon passes a direct cooling zone, the third heat exchanger, and reaches the top gas distributor of the adsorption zone via pneumatic conveying. Liquid phase adsorption is used mainly to bind turbid substances, to purify and decolor solutions, and to separate organic pollutants. Grained active carbon is employed as an adsorbent, mixed in a pulverized form with the liquid or as particles in a percolation process. A schematic for wastewater treatment by adsorption of waste substances, for example, is shown in Fig. 4-5. The adsorption process is carried out discontinuously with grained active carbon or continuously with grained or pulverized active carbon. Figure 4-6 shows a discontinuously operated adsorption unit including three adsorber vessels connected in series, and filled with active carbon beds for the pretreatment of wastewater. Regeneration

Lurgi KONTISORBON process. Representation according to data of Lurgi, FrankfurUMain I4.15, 4.161.

4 Fig. 4-4.

Cooling water

1 x 1 Heating medium

a

Nitrogen Nitrogen - solvent mixture Active carbon

nAir r T d Solvent Process data:

287

Superficial gas velocity of fluidized bed: 0.8-1.2 m/s Pressure drop (total): 10-18 mbar (ca. 25-5070 of pressure drop of packed bed units) Height of each fluidized bed: 0.025-0.05 m Adsorbent: Active carbon, particle diameter 0.7 mm Active carbon losses: ca. 2%/a (referred to the circulating carbon mass) Inert gas feed: ca. 0.05% of clean exit air Regeneration energy consumption: 30-4070 of a packed bed heat requirement

288

4 Adsorption

0)

W a s t e w a t e r composition

I

I

f

substances

Settling substances

Floating substances

I I

I

z-

Suspended substances

substance

I

substance

Nonbiode-

substances

__ I 1 /Sedimentation

i

Filtro tion

1

1

i

substances

b

Wastewater gathering points 1-5

Adsorption

c

Adsorption

-

Biology

Adsorption -

+Active c a r b o n (pulverized) C

Fig. 4-5. Systematization of wastewater composition (a) and adsorptive wastewater treatment (b). Representation according to UHDE [4.17]. a) Partial flow treatment b) Adsorption of toxic components prior to biological treatment c) Continous feed of active carbon into the activated-sludge tank d) Further treatment of persistent (e. g., biocide) wastewater components

4.1 Principles of Adsorption and Desorption, Processes and Examples

289

Fig. 4-6. Discontinuous adsorption unit for pretreatment of wastewater. Representation according to UHDE [4.17, 4.191. 1 Adsorber vessel with active carbon packing 2 Condenser 3 Gravity separator

or reactivation of active carbon is carried out by steam or outside in an reactivation oven. Figure 4-7 shows a continuously operated unit for the downstream cleaning of pretreated wastewater. Biologically treated wastewater enters the countercurrent adsorber from the bottom via a distribution system, and leaves by means of a chute system. The upflow velocity is ca. 6.5 m/h. Active carbon continuously slides through the adsorber and is discharged by an air lift pump. Transport water and carbon are separated by a sieve. The carbon is then dosed by a screw feeder into a two stage rectangular fluidized bed oven employed as a reactivation oven. With a flue gas temperature range from 20O-105O0C, active carbon is reactivated and withdrawn glowing. After quenching in a water bath, the active carbon is transported by water back to the adsorber. Flue gas loaded with expelled organic substances from the reactivation oven

are treated in an secondary-burning treatment unit. Table 4-3 gives an overview of technical application of adsorption used for gas drying, separation of gas mixtures, gas generation, gas enrichment, gas cleaning, and wastewater treatment by means of selected examples. Table 4-4 gives the effective pore size of different molecular sieves and adsorbate molecule diameters. The following advantages characterize adsorption as a thermal separation process : 0

0

0

It may be carried out under ambient temperature and hence under favorable energetic conditions. The process is positively influenced by special treatment of the adsorbents It is possible to treat mixtures with low adsorbate concentrations if a suitable adsorbent is available Simultaneous separation of several mixture components is possible

290

4 Adsorption Exit water (clean wastewater)

I

-

Active carbon reflux

-

I

Flue gas t o chimney

Pretreated wastewater Gas or oil

Fig. 4-7. Adsorption unit for continuous treatment of wastewater. Representation according to UHDE[4.17, 4.181. 1 Countercurrent adsorber 2 Air-lift pump 3 Drain screen 4 Active carbon feed 5 Desorption and reactivation in a two stage fluidized-bed furnace 6 Quench vessel 7 Bin for active carbon make-up 8 Postcombustion unit

Table 4-1. Actual data of a plant to separate dioxane*. Process parameter Wastewater input Loading rate Contact time (3 adsorber) Dioxane feed concentration Dioxane exit concentration Degree of reduction Adsorption capacity (referred to dioxane and active carbon) Holdup time Stripping time Steam consumption (referred to active carbon)

*

(m3/h) (m3/(m2 . h)) (h) (mg/l) (mg/ 1) (070)

Representation according to data of UHDE [4.17, 4.191.

Design

Operating state

16 3 2.5 100

25 4.7 I .6 16 1

2

98

94

4.2 Adsorbents, Selection of Adsorbent

291

Table 4-2. Actual data of a plant for subsequent treatment of 3000 m3/d wastewater*.

Loading before adsorption (mg/l)

Loading after adsorption (mg/l)

BOD5

BOD5 COD

COD TOC

20 50-80 20-60

TOC

5

1-27, mean 7 2-30. mean 8

BOD 5 Biochemical oxygen demand within 5 days Chemical oxygen demand Total organic carbon

COD TOC

* Representation according to UHDE [4.17, 4.191.

4.2 Adsorbents, Selection of Adsorbent 4.2.1 Adsorbents Adsorbents are solids characterized by a high specific mass or volume surface area. There are a large number of homogeneous and heterogeneous active centers on the surface where free bonds are provided by the solid for the adsorption of adsorbate molecules. The high specific surface is mainly due to the internal surface of the adsorbent body, which consists of a macro-pore system and a particularly important micropore system with external excess. Pore sizes are commonly : 0

0 0 0

Macropores with a diameter dpr > 50 nm Mesopores with 50 > d,, > 2 nm Micropores with 2 > dpr > 0.4 nm Submicropores with dpr < 0.4 nm

The pore size distribution characterizes the pore system and hence the specific surface and the adsorbate loading capacity of the adsorbent (measurement techniques to de-

tect micro- and macropore systems are found in [4.20, 4.261). Theselectivity of the adsorbent and hence the possibility of gas mixture separation by adsorption, is based on three effects, the equilibrium effect, the steric effect, and the kinetic effect. The equilibrium effect causes the key component to be bonded more strongly to the adsorbent than the others. With the steric effect, micropores of the adsorbent only allow molecules with a diameter smaller than the pore sizes to pass (sieve effect with adsorbents of uniform and very narrow micropore size distribution). In the case of the kinetic effect, key component molecules diffuse faster into the pore system of the adsorbent than less movable and slow molecules [4.11].

4.2.2 Requirements for the Adsorbent, Adsorbent Selection To ensure economic, safe, and nonpolluting operation, the adsorption process has to meet certain requirements. These requirements, which at the same time are aspects of adsorbent selection, are listed in Table 4-5.

292

4 Adsorption

Table 4-3. Adsorption application (selected examples). Process

Adsorbent

Air and gas drying

AA, SG, Zeolith-MS

Gas production, gas mixture separation, gas enrichment H2 N2 0 2 0 3

nho-Separation from hydrocarbon mixtures (see Table 4-4)

Gas purfication Solvent recovery

Production stage Ad De

+

Remarks Process Company PSA, “Heatless Drying” PSA

Zeolite-MS/AT C-MS C-MS Zeolite-MS C-MS Zeolite-MS/SG Zeolite-MS

AC AC SG

+

+ +

+

+

+

+ +

+

+ +

UCC/HYSIV Linde/UCC BF-N2-process Linde/BF-process BF-02-process Linde-Ozon-process UCC-IsoSiv-process

Supersorbon-process (Lurgi) Kontisorbon-process (Lurgi)

AC

+

ADSOX-process

Separation of SO2

AC, MS ACO

+ +

Babcock-BF-process

Separation of NO, and SO2 Separation of H,S, CS2 Separation of H2S

ACO AC, MS, AA Fe(OH),-rich adsorbent (hydrated iron oxide etc.) AC, AA, Biofilter

Solvent enrichment with subsequent oxidation Flue gas treatment

Deodorization

Wastewater ireatment (Separation of turbid matter, organic matter, etc.) ~~

AC

+ + +

BF-Uhde-process Giulini-Saar-Ferngasprocess

+ (if possible)

~

Abbreviations : Ad Adsorption; De Desorption; AC Active carbon; AA Active alumina; ACO Active coke; C Carbon; MS Molecular sieve; SG Silica gel; PSA Pressure swing adsorption

4.3 Adsorption Kinetics

293

Table 4-4. Effective pore diameter of different molecular sieves and corresponding adsorbate with smaller critical molecular diameter [4.1 I]. Effective pore diameter (lo-'' m)

Adsorbate

3-3.8

He, Ne, Ar, CO, NH,, H2, 02,N,, H20, CO,

4.2

Kr, CH,, C2H6, C2H4, C2H2, SO2, CH30H, CpH50H

5 .O

n-C3H, to C,,H3,, C2H5C1,Cyclopropane

8.0 9.0

i-C4Hlnto i-C14H3n, Dioxane, Cyclohexane 1.3,s-Triethylbenzene

Table 4-5. Adsorbent requirements and selection aspects.

0 0 0 0 0 0

0 0

Even at low adsorbate concentration in the gas or liquid phase high additional load of absorbed material (additional load is the adsorbent load with adsorbate when the adsorption process is finished minus the residual load after desorption). Large useful capacity. High selectivity to certain adsorbates. Good desorption behavior, easy to regenerate, low residual load of the key component. High adsorption rate, short adsorption or mass transfer zone, MTZ. High stability to heat and vapor. Good chemical restistance, low ignitability. High abrasion resistance. Low specific packing pressure drop. Low price

4.2.3 Technical Adsorbents, Characteristic Data of Adsorbents Technical adsorbents include active carbon, active coke, silica gel, active argillaceous earth or aluminum oxide gel, bleaching earth and molecular sieves (Table 4-6). In Tables 4-7 and 4-8, some properties of the adsorbents are described.

ity Xst,see Chapter 1.4.4.2) and adsorption kinetics. The adsorption rate is described as a function of the limiting transport phenomena for the adsorbate transfer from the bulk of the carrier phase to the inner of the adsorbent particle. Individual transport steps are (Fig. 4-9): 0

0

4.3 Adsorption Kinetics The design of adsorbers depends on the adsorption equilibrium (static loading capac-

0

0

Free diffusion out of the carrier phase into the boundary layer around the adsorbent particle (exterior mass transfer) Diffusion through the boundary layer (boundary layer diffusion, outer diffusion) Diffusion inside the pore system of the adsorbent Addition at the inner pore surface (actual adsorption)

294

4 Adsorption

Table 4-6. Applied adsorbents. Adsorbent based on carbon [4.4, 4.20, 4.561 Active carbon, active coke, carbon molecular sieve, carbon fiber mat, carbon fiber paper [4.54, 4.551. 0 Carbon content >go%, others are inorganic salt and ash. 0 Base material and manufacturing: wood, brown coal, pit coal, peat, saw dust, coconut shells, petrol coke. Carbonization including following pore structure increased by steam activation or chemical activation using phosphoric acid solution or zinc chloride solution, if necessary oxidation using air or steam. 0 Amorphous solid structure, graphite lattice of microcrystallite, powder coal, granular coal of particle size ca. 0.25-4 mm. 0 Hydrophobic nature, especially suitable to bind nonpolar substances, flue gas cleaning, gas cleaning, solvent recovery, deodorization, respiration filter, catalyst carrier, separation of gas mixtures, deoiling, degreasing, liquid decoloration; organic substance removal from drinking water, feed water, and wastewater; chlorine and ozone removal from water.

Active argillaceous earth, aluminum oxide gel, alumina gel 0 0 0

> 8 5 % A1,0,, others are inorganic salt. Surface active aluminum oxide, activated by calcination, pellets or extrudate, particle size ca. 2-8 mm. Suitable for gas drying, gas enrichment, polar substance adsorption from solutions, catalyst carrier

Silica gel [4.21, 4.221 0

0 0

Granulated, porous amorphous form of silica dioxide, produced with sulfuric acid and sodium silicate hydrate, >99% Si02,others are inorganic oxides, dry pellets (size 3.5 mm) contain ca. 3 % A1203. Narrow pores and wide pores, pellet size ca. 2-8 mm. Suitable for gas drying and catalyst carrier.

Molecular sieve-zeolite (MS) [4.23, 4.241 0

0 0

0

0

Natural or synthetic tecto silicate with - a three dimensional aluminum silicate crystal lattice consisting of S i 0 4 and AIO, tetrahedron - a system of cavities inside the lattice connected by pores of defined and absolute equal pore diameter - free movable cations, exchangeable in a solution, to compensate the negative charge of the acid ion lattice - the general formula x[(M',M",,~) * A102]ySi0, . zHzO (MI, M" monovalent cation and divalent cation, i. e., Na or Ca) for example: Na12(A102)12(Si02)12 . z H 2 0 MS 41 nm Preparation by crystallization of gels generated in aqueous alkaline silicate- o r aluminatesolution, also by adding amine or ammonium compounds. Large adsorption area (800- 1000 m2/g), high electrostatic adsorption force, feed pores of defined diameter classify different molecule sizes (screening effect), pore diameter according to MS-type 30- ca. 100 nm. Delivery forms: powder, pellets, particle size ca. 0.3-5 mm. Suitable for gas drying, gas cleaning and gas mixture separation, as well as catalysis purposes.

295

4.3 Adsorption Kinetics

Table 4-7. Adsorbent characteristics for evaluation. Specific surface area a: outer and inner adsorbent surface area referred to its mass unit (m2/g). Experimentally determined by N2-adsorption isotherms with respect to space requirement of an N2-molecule (BET-surface area). Density, adsorbent porosity, packed bed porosity Solid density (true density, skeleton density) Q ,

m,

Skeleton mass, V, Skeleton volume

Particle density (apparent density)

e mL

m

v

= = I

m,+m,

K+G

dry adsorbent

-- m, v,+G

(4-2)

Air mass in the pore system, Vp Pore volume

Particle density

Q

~ of ,adsorbent ~

ep,x = ep * (1

X

eP of

plus adsorbed component

+ x)

(4-3)

Adsorbent load with adsorbed component (dimensionless)

Adsorbent porosity cp

= -m, =

m,

v, V . ( l - E p ) e p = e, . (1 - Ep) @

=

Packed bed density (bulk density) ing volumeter)

mb

v,

&

m V.(l-Ep)

=-1

ep

-Ep

n (4-5)

@ b (adsorbent

bulk density after tapping using a stamp-

Packing mass (for example adsorbent packed bed mass) Packing volume Packing porosity (packed bed voidage) (continued next page)

296

4 Adsorption

Table 4-7. (continued) 0

Approximate values of density and porosity Adsorbent

@$

@p

( W ) (Wl) ~

Active carbon Active carbon, fine-pored Active coke Silica gel, fine-pored Silica gel, wide-pored Active argillaceous earth Molecular sieve

2.2 2.0 1.9 2.2 2.2 3.0 2.6

@b

(Wl)

EP

-

-

~~

~~~

0.6 0.8 0.9 1.1 1.1 1.2 ca. 1.3

ca. ca. ca. ca. ca. ca. ca.

0.4 0.45 0.6 0.75 0.60 0.75 0.75

0.73 0.60 0.53 0.50 0.50 0.60 0.50

0.33 0.44 0.33 0.32 0.45 0.38 0.42

Pore size distribution of the micropore system determined by the water adsorption isotherm using the Gibbs-Kelvin equation; determined for the macropore system by mercury penetration and mercury or helium displacement, an example is given in Fig. 4-8. Fig. 4-8. Differential pore size distribution of a Linde molecular sieve and molecular sieve carbon [4.27]. W

L

0

Pore radius-

0

Adsorption capacity X (adsorbed component capacity of the adsorbent)

mi

ST

Maximum mass of adsorbed component Adsorbent mass

The static adsorption capacity, “static activity” X,,may be derived from the respective adsorption isotherm (see Chapter 1.4.4.2). A steep course of the isotherm in the log XJlog p i diagram gives a favorable adsorption. Additional adsorbent consumption has to be taken into account during design to consider adsorbent ageing, adsorbent damage, adsorption displacement etc. [4.28, 4.29, 4.341.

Catalytic properties If an adsorbate has to be converted into a chemical compound, i. e., separation of SO2 or

H2S from flue gas, a catalytic activity of the adsorbent is desired. Due to the catalytic activity of active carbon the following reactions are favored: SO2 + H2O + 1/202 + H2SO4 H,S+ 1/20, + H , O + S SO2+2H,S - j 2 H 2 O + 3 S If the adsorbate has to be recovered, no adsorbent catalytic activity is allowed.

Adsorbent properties [4.9, 4.10, 4.13, 4.25, 4.271.

500-1000

0.25-0.30 0.30-0.40

100-400

0.40 0.10

700-850

250-350

0.30-0.45 0.05-0.10

400- 800

600-850

0.35-0.45 0.10

700-800

100

0.05-0.10 0.20-0.30

600

e

600- 1000

0.25 -0.40 0.40-0.50

400-500

on

1000-1500

0.30-0.50 0.50- 1.10

300-500

on

(cm3/g)

(kg/m3)

Pore volume

Bulk density

s sieve 600-900

- Micropores - Macropores

(nm)

(m2/g)

Mean pore diameter

Specific surface

20

Specific heat (kJ/(kg. K))

Thermal conductivity (W/(m. K))

Desorp temper ("C) 100-1

0.84 0.84

100- 1

0.84 21 100

30-35

3-10

0.92 0.92 0.85-1.05

0.95- 1.05

0.20 0.20 0.12

0.13

120-2 120-2 150-3

200- 3

Boundary layer

.........

Pore diffusion

Surface diffusion

........

?[ .... ....

dpr<0.L nm

. . * I

Active crack or edge diffusion Boundory layer (fluid film)

Free diffusion Film diffusion Solid diffusion

I diffusionccsolid diffusion

.

.

I t -

r

+ diffusion<
Fig. 4-9. Adsorbent structure (a), pore system (b) and adsorbate diffusion (c and d). Representation according to [4.35]and [4.51].

.....

.:.: ... .... .....

299

4.3 Adsorption Kinetics

The slowest mass transfer step controls the total adsorption rate. With low carrier phase flow velocities, a pronounced boundary layer develops around the adsorbent particles; the diffusion resistance of the boundary layer is larger than that inside the particles; the outer diffusion limits the pro-

cess (often the case with liquid phase adsorption). With higher carrier phase flow velocities, the internal diffusion of the particles limits the adsorption rate. It mainly depends on the pore structure of adsorbent which allows different types of solute diffusion mechanism (Table 4-9).

Table 4-9. Mass transfer during adsorption. Particle diffusion.

Diffusion mechanism 0 In macropore system free diffusion ( A s rpr) 0 Pore diffusion (Knudsen diffusion A = rpr) 0 Micropore or crack diffusion (active crack diffusion r, = rpr) 0 Diffusion on the pore surface in the adsorbed phase [4.31]

Total particle diffusion [4.9,4.251 : (4-9) Approximation (at linear adsorptions isotherm course, heat effects are neglected and at constant adsorbent surface partial pressure of the adsorbate) (4-10) Mean free path of adsorbate molecules Pore radius rA4 Molecule radius k Adsorbate transfer coefficient dr; Diffusion length X,,R(t),X, Initial adsorbate load, mean adsorbate load at time t and final adsorbate load of adsorbent particle t Time @P Adsorbent particle density Deff Effective diffusion coefficient A

'pr

Mean diffusion coefficient of steam, D e f f at , 15-20°C [4.32, 4.331: Adsorbent

Deff(lo-'' m2/s)

Active carbon Active argillaceous earth Molecular sieve 4 nm Silica gel

0.8 2-4 1-2 0.5- 1.5

300

4 Adsorption

The effect of mass transfer on the chronological course of the adsorption process in a technical adsorber with fixed bed is described by the mass transfer zone MTZ and the breakthrough curve. In the mass transfer zone (Fig. 4-10) transport of adsorbate takes place from the carrier phase to the adsorbent. Adsorbate loading of the carrier phase starts with an initial concentration Y, and decreases to the desired final concentration Y,. Correspondingly the adsorbent concentration increases from X, = 0 (fresh adsorbent) or X, = Xreg(regenerated adsor-

bent), to the maximum loading capacity X,. With an increasing contact time c or increasing quantity G of the carrier phase, the mass transfer zone moves in the direction of loading through the adsorber bed until the end is reached and the adsorbate finally breaks through, the adsorbate concentration in the exit increases (Fig. 4-11). Breakthrough curves for an adsorber are found experimentally. The adsorbate concentration Y, of the carrier phase is measured at the exit of the adsorber or after the adsorber bed and plotted against time t or quantity G . It is often found to be related

x.

-0

A

MT2

Fig, 4-10. Mass transfer zone MTZ and length of unused bed LUB. X, Y Adsorbent load, carrier phase load (with key component or adsorbate) X, Dynamic equilibrium adsorbent load X, Initial adsorbent load (after regeneration) z Packing height

4.4 Variations of Adsorption, Design of Adsorbers

to the entry concentration. Figure 4-11 shows the development of a breakthrough curve and also distinguishes between breakthrough curves for adsorption systems with “good” or “poor” kinetics. With good kinetics, the rate of the adsorption process is so fast that initially no adsorbate concentration Y,,, is detected at the adsorber exit. After some time, a rapid increase in the concentration is observed and so the breakthrough curve shows a steep gradient. With poor kinetics and a low adsorption rate, the adsorbate concentration Y, increases gradually with loading time and slowly reaches the final value. Generally, small breakthrough concentrations are required (mean concentrations over the complete adsorber packing cross section after the breakthrough). Thus good kinetics results in a good adsorber efficiency, and poor kinetics causes a low utilization of the adsorber. With coadsorption of two or more adsorbates, each component has its individual breakthrough curve, but due to the mutual influence its form is different to that for single component adsorption. At the start of the adsorption process, the loading of the adsorbent corresponds to the diffusion velocity and the concentration of the adsorbate in the carrier phase. If the component A is more favorably absorbed it (Fig. 4-12) displaces the other component B from the inner adsorbent surface. The loading capacity of A is barely reduced by the presence of B. However, A forces a strong decrease in the loading capacity of B. An earlier breakthrough of B results from the presence of A and the concentration of B in certain layers is larger than the initial concentration (“overshooting”). This is due to the displacement of adsorbed B by the moving adsorption wave of A. Breakthrough curves, particularly for coadsorption, should be found experimentally under conditions as close as possible to the operating conditions. Methods to cal-

301

culate breakthrough curves are described in the literature (e. g., [4.9,4.30], and for coadsorption [4.31, 4.321). These methods are based on restricting assumptions such as exclusion of any coadsorption, steady-state adsorption process, special course of the adsorption isotherm, no heat effects, etc. Unfortunately, the assumptions are only relevant in a few rare practical cases.

4.4 Variations of Adsorption, Design of Adsorbers With a favorable equilibrium and a high adsorption rate, the adsorption may be carried out in a single stage adsorber with packing. If, in a single stage, the desired concentration is not reached, two or more adsorption stages are connected in series. This is realized by cross or countercurrent flow of the adsorbent and the fed carrier phase.

4.4.1 Single Stage Adsorption in a Vessel Adsorber with Adsorbent Packing With adsorption in a fixed bed an adsorbate-rich carrier fluid, for example a gas stream G,, flows through an adsorbent over the loading time tg.The concentration of the adsorbate component is reduced from the feed concentration Y, to a maximum allowable exit concentration Y, by adsorption at the adsorbent surface in the adsorption or mass transfer zone, MTZ. At the start of the loading phase, (t = 0) the MTZ starts at z = 0. With increasing loading time, the MTZ moves in the loading direction away from the gas entry through the adsorbent fixed bed (t,, t2, . . ., Fig. 4-11). Under the MTZ (with a loading direction

x, YL-

bl

t

Y

fl

C) 4

b '

P

zh

f-

-I

-I

'b

'h

/

0.5

I

'b,P

'b.G

'h

t-

Fig. 4-11. Movement of the mass transfer zone in a packed bed adsorber (a) and breakthrough curve (b). Breakthrough curves for adsorption systems with good and poor kinetics (c). FA Fresh or regenerated adsorbent zone MTZ Mass transfer zone LAZ Loaded adsorbent zone BC Breakthrough curve (G good kinetics, P poor kinetics) Carrier phase load (with key component or adsorbate) Y t Time

4.4 Variations of Adsorption, Design of Adsorbers

303

where @b is the bulk density and AQ the cross section of the adsorbent bed. AQis determined by means of a gas load factor F F = wg-l/e,-O.2

or

f-

Fig. 4-12. Binary breakthrough curves for the coadsorption of two components A and B. BC A, BC B breakthrough curves of A and B

from bottom to top) the adsorbent is saturated with adsorbate or adsorbed substance (zone LAZ in Fig. 4-11, saturation concentration X,). Fresh or regenerated adsorbent remains above the MTZ with the initial absorbate concentration, X,. This region has not yet taken part in the adsorption process. If the front of the MTZ reaches the top of the packing, a breakthrough of adsorbate occurs; the loading phase ends and the adsorbent has to be regenerated. During the loading period, tg, the effective adsorbent mass S , of the bed adsorbs m, kg of adsorbate

]/pa

... 0.4-* S

(4-13)

where wg is the superficial velocity of the gas mixture of density eg related to the empty adsorber cross section. As seen in Fig. 4-11 at the end of the loading period tb a breakthrough of adsorbate begins. In the mass transfer zone, the adsorption capacity is not fully exploited. The MTZ may therefore be considered as consisting of a saturated zone, where the adsorbate concentration X , is already reached, and an unloaded zone (the length of unused bed LUB), (LUB model 14.28, 4.361). If a symmetric s-shape course of the loading curve X ( z )is assumed, the adsorption capacity of the MTZ is only 50% utilized, i. e., LUB

1

- . MTZ 2

(4-14)

With an assumption of constant expansion of MTZ and a constant moving velocity wMTZ of MTZ, it is

m i= G,. ( Y , - yW) - tg= S , - ( X , - X,)

(4-11)

where G, is the inert carrier fluid flow rate. The height of the packing Z in the adsorbent bed is fixed by S,

where time until increase in load is Y/ Y, = 0.5 at the end of the packed bed time until the start of breakthrough above Y, (Fig. 4-11) height of fixed bed (height of adsorber packing)

304

4 Adsorption

The periods th and t b to determine wMTZ and LUB are found by experiment. For any loading profile X ( t ) , MTZ and LUB are found according to Table 4-10. Table 4-10 also illustrates the equation system to calculate the adsorbate concentration c, at any time and location in a fixed-bed adsorber including a numerical approximation by ROSEN[4.30]. The required height of the adsorbent packing Z (Eq. (4-12)) is sufficient to adsorb m ikg adsorbate if the time t b until the beginning of breakthrough is chosen as loading time tE. If the loading period is fixed as tg = th, the available effective height of the bed is then

z,,=,z,- L u B

1

z h - - . MTZ

2

(4-39)

The energy of conveying is essentially determined by the pressure drop A p of the carrier phase flowing through a fixed-bed adsorber of height z,and is, for example, calculated by means of the ERGUNequation [4.37] k , (1 - c ) ~qg wg + -Ap _ *

z

*

*

c 3 * d:

+

k, * (1 - E )

*

eg

*

Wf

(4-31) 1

e 3 * dp

-

I

specific constants of the packing (often, k , = 150, k, = 1.75 and for Baylith molecular sieves, for example k , = 200 and k, = 1.75) apparent void fraction of the packing (Table 4-7) dynamic viscosity (Pa s) and density (kg/m3) of the carrier phase superficial gas velocity related to the cross section of the packing A, (“empty tube velocity”) diameter of adsorbent particle

-

An enthalpy balance over the fixed-bed adsorber, according to Fig. 4-11 gives

With adiabatic operation, Q = 0, the temperature of the adsorbent bed starts from an initial value tYS,, at the beginning of the loading phase and increases to the final value ds,S,W. Since the heat of adsorption QAd

ST *

(xu- xu)hAd,d s, *

*

AhAd (4-33)

mainly remains in the adsorber bed, a decrease of the loading capacity X , results from the temperature increase, which has t o be considered when the quantity of mass of adsorbent ST required is calculated. (A reliable design basis, considers X, at tY,,). In Eqs. (4-32)-(4-33) the nomenclature used is enthalpy of the entry and exit carrier mass flow rates G,, G,, respectively A,,,, h , , enthalpy of the adsorbent (without loading) at the start and end of the loading phase AhAd, hA,d integral and differential adsorption enthalpies

hg,a,h , ,

The procedures for differential adiabatic adsorption in a fixed bed, and adsorption in an adsorber with jacket cooling, can be found in [4.38, 4.391 and the modeling of an isothermal fixed-bed adsorber in [4.50]. With isothermal operation of the adsorber, the bed temperature d, is kept constant, and the heat (2 to be removed during the loading phase is calculated by Eq. (4-32). The adsorption ,enthalpy AhAd is the sum of the adsorbate condensation enthalpy and the binding energy Ahb

4.4 Variations of Adsorption, Design of Adsorbers

305

Table 4-10. Mass and heat transfer in an adsorbent packed bed. Correlation between LUB and MTZ (Fig. 4-10) ZE

pw - x t , dz (4-16)

KMTz

Mass transfer zone form factor

Mass and heat transfer in an adsorbent packed bed 0

Outside (bulk) mass and heat transfer Adsorbate mass flux to adsorbent surface mi = /3. (cFK- cFG) Heat flux 4

c, fiK,

=a

. (L9K - L ~ G )

(4-17) (4-18)

cFG Adsorbate concentration in gas kernel and adsorbent surface L9G Temperature in gas kernel and adsorbent surface

Correlations by ACETISand 'I~ODOS [4.53] to calculate steady-state mass and heat transfer coefficients in packed beds. NU =

Sh

Re, Pr Nu, Sh

=

1.1 ' Re. Pr1I3 - 1.5

(4-19)

0.725 . Re. Pr ' I 3 - 1.5

(4-20)

Reynolds and Prandtl nos. of gas phase Nusselt and Sherwood numbers (4-21)

a, p

Heat and mass transfer coefficients Adsorbent particle diameter I,, Dg,vg Thermal conductivity, diffusion coefficient and kinematic viscosity of the gas phase

dP

0

Internal mass transfer (see Table 4-9)

Equation system to describe the concentration profile in an adsorbent packed bed 8cF + (1 - E ) . aiat., + E . =0 at

0

acF Adsorbate mass flux balance w, __

0

External mass transfer

(4-23)

Particle diffusion

(4-24)

az

-

(4-22)

(continued next page)

306

4 Adsorption

Table 4-10. (continued)

0

Approximation by ROSEN [4.30] Bed height parameter A

=

Contact time parameter B

(3A /2B) - 1 2.1/(1

CFLl

+ 5c)/5B

1 2 . D e f f * K * * z *-(E1) wg * d j =

Mass transfer parameter C =

8 . Deff. ( t - z * E / w ~ )

4

2 * Deff K*

(4-26) (4-27) (4-28)

dp . P

where: cF, cFG

c,, csG erf ( x )

Adsorbate concentration in the gas and at adsorbent particle surface Adsorbate concentration inside adsorbent particle and on the surface; at linear adsorption isotherm with slope K * it is c,, = K * . cFG Error function (4-29)

r, dp E

Deff WE

hhAd = Ah,,

Adsorbent particle radius and diameter Adsorbent packed bed void fraction Effective pore diffusion coefficient Superficial gas velocity

+ hhb

(4-34)

jiAd,d may be calculated from the course of two adsorption isotherms (see Chapter The differential adsorption enthalpy hAd,d, 1.4.4.2) given in kJ/kg adsorbed substance, is exchanged when the adsorbent load is increased from X to X dx.The integral adsorption enthalpy AhAd, given in kJ/kg adsorbent, is the evolution of heat if 1 kg ad(4-36) sorbent is loaded with adsorbate, starting from 0 to the final load X , or upon integration

+

The differential adsorption enthalpy hAd,d depends on the adsorption system, the adsorption conditions, pressure and temperature, and the adsorbate concentration X . Its value decreases with increasing X .

Figure 4-13 gives the differential molar adsorption enthalpy for the adsorption of different hydrocarbons on active carbon.

4.4 Variations of Adsorption, Design of Adsorbers

307

to the desired final concentration Y,. The cross flow of carrier and adsorbent phases in this manner is favored if the adsorbate has to be largely removed from the carrier phase. Conversely, the loaded adsorbent may be regenerated by stagewise contact with a regeneration fluid in a cross flow cascade. An adsorbate balance over stage 1 using the notation in Fig. 4-14, gives

-' 5

293 K

Fig. 4-13. Differential molar heat of adsorption

adsorption system hydrocarbons/activated carbon) [4.40]. T, Boiling temperature X = 0.03 kg/kg (adsorbed component/active carbon), adsorption temperature 20 "C hAd,d,

If the adsorption equilibrium of phases leaving the stage is just reached, the result of stage 1 is characterized by the loading differences y, - r, and X, - X , in the operating diagram, Fig. 4-15 (see also Chapter 1.10). The slope of the operating line for stage 1 is tan

4.4.2 Multistage Adsorption with Cross Flow of Gas and Adsorbent Phases

3tl =

-

y*-

r,

Xl -Xu

=

s, -~

(4-39)

GT

This is fixed by the adsorbent mass flow rate and corresponds to the adsorbent ratio v, of stage 1. This can be analogously applied for stages 2, 3, . . ., N,in which the In cross flow multistage adsorption, the desired exit concentration Y, of stage N is carrier phase flows according to Fig. 4-14 in either just reached or lower. turn through adsorber stages 1, 2, .. ., N. If the adsorption equilibrium in each Each adsorber stage is supplied with adsor- stage is not reached, the number of practibent with the same initial concentration cal stages Np for the cross flow cascade is X u . The adsorbate concentration in the calculated from N by consideration of an carrier Y, decreases stagewise to q, y2, . . . empirically determined stage efficiency.

Fig. 4-14. Multistage cross-flow adsorption.

$,

308

4 Adsorption

x,

x-

Fig. 4-15. Operating diagram of a three stage cross-flow adsorption. OL Operating line EC Equilibrium curve Adsorbate load in the carrier phase Y Adsorbed component load of adsorbent X

4.4.3 Multistage Countercurrent Adsorption With multistage countercurrent adsorption, the receiving adsorbent phase passes from top to bottom in an adsorption column, while the carrier phase flows in the reverse direction toward the adsorbent phase (Fig. 4-16), i.e., upward. If only one adsorbate component is removed from the carrier phase, the determination of the number of theoretical separation stages Nt of the adsorber and the calculation of the active adsorber height for mass transfer 2 are analogous to countercurrent absorption. With the notation in Fig. 4-16 the mass of adsorbent ST required is (4-40)

Fig. 4-16. Multistage countercurrent adsorption. 1,2,. . ,,N Separation stages y Adsorbent load and carrier load with adsorbate or key component

x,

An adsorbate balance over the top section of the adsorption column gives

Y = 7X . + Y, - 7X. , ST

ST

GT

GT

(4-41)

4.4 Variations of Adsorption, Design of Adsorbers

the operating line in the loading or operating diagram for countercurrent adsorption shown in Fig. 4-17. From this diagram, the required number of theoretical separation stages, Nt is found by drawing steps between the equilibrium curve and the operating line. Also, Xu,,,,and therefore, the

309

Following the HTU-NTU concept (see Chapter 1.9.4) the height of the adsorption column Z is calculated, for example, by

Z=

minimum value vmin of the adsorbent ratio, v = S T / G T can be obtained from Fig. 4-17.

GT . --dY AQ-itg-a, Y - Y*

= HTUo,g

*

NTUo,g

(4-42)

Fig. 4-17. Determination of the number of theoretical stages in countercurrent adsorption at minimum adsorbent ratio vmin and respectively. EC Equilibrium curve OL Operating or balance line for v OLM Operating or balance line for vmin

t

Y

310

4 Adsorption

where AQis the cross-sectional area of the adsorption column, a, the volume specific surface, kg the overall mass transfer coefficient related to the gas phase and Y * the equilibrium concentration corresponding to In contrast to absorption HTUvalues for adsorption are mainly found experimentally as there is little published literature. The design engineer is advised to apply the experience of companies that manufacture adsorbents, and use their facilities to carry out experiments at pilot plant scale under conditions as close as possible to the operating conditions. If several adsorbate components are adsorbed to a different extent by the adsorbent, counterflow fractioning is required for the separation. For example, in a rectisorption or hypersorption process (Fig. 4-18), a gas

x.

Feed hopper

.!

Heating{

4.5 Adsorber Types The adsorber type is essentially governed by the chosen operating conditions. Discontinuous adsorption operations distinguish themselves by: 0

Light compon

Heavy component

mixture flows in an adsorption column toward the downward-moving adsorbent. In the bottom section of the column, the loaded adsorbent is regenerated by heating with steam. Components with a low volatility in the gas mixture are released and withdrawn at the bottom. Highly volatile components leave the column at the top. The regenerated adsorbent is transferred using pneumatic conveying to a storage bunker, cooled, and fed back to the column top or the loading zone. The rectisorption process is applied to separate hydrocarbon mixtures, particularly to separate valuable gas components from a gas mixture, for example, ethylene from coke oven gas.

zone

1S;;itionating

I

~--Desorption zone Adsorbent discharge grid

0

Disadvantages are: 0

0

0

Pneumatic conveying

Circulation blower

Fig. 4-18. Hypersorption process for continuous adsorption. Representation according to WIRTH[O.l, Vol. 21.

Simple adsorber construction and experienced technology Little mechanical and thermal wear on the adsorbent

0

Large amounts of the adsorbent and of desorption auxiliary substances are needed Limited adsorber and desorber dimensions and if necessary parallel operation Sensitivity to dust deposits, poisoning and thermal stress in the adsorbent High expense for process control facilities

Discontinuous adsorption is mainly carried out in a vessel absorber with an adsorbent fixed bed, an adsorbent of honeycomb structure, or an adsorbent ring layer, and sometimes in a one stage fluidized bed

4.6 Desorption, Regeneration of Loaded Adsorbent

(Figs. 4-19 and 4-20). An even flow and the best possible distribution of the fluid phase in the adsorbent packing has to be ensured. Advantages of continuous adsorption operation are : 0

0

0

0

Smaller adsorber and desorber dimensions, with simple process control equipment Shorter desorption phase at a higher temperature level Less adsorbent and desorption auxiliary substance are needed Relatively insensitive to dust deposits and poisoning of the adsorbent

Disadvantages are the high expense on peripheral conveying equipment and the higher attrition of the adsorbent. Continuous adsorption mainly takes place in single or multistage fluidized beds (Figs. 4-4 and 4-20b). Small mesh sieves are used with an adsorbent downcomer or large mesh sieves with the adsorbent raining through the holes (for more information on fluid dynamics see [4.41, 4.421, on residence time [4.43], and on pressure drop, flow through particles, particle concentration [4.44-4.471).

3 11

Continuous adsorption may be also carried out in moving fluidized channels and moving beds (Fig. 4-20). In the case of a fluidized trough, a fluidized adsorbent mass with a small bulk density is forced to move in a guided direction from zones of higher bed height to zones with lower bed height. In a moving bed adsorber, [4.48] the slightly loosed up adsorbent moves cross flow or countercurrent to the fluid phase.

4.6 Desorption, Regeneration of Loaded Adsorbent Whereas adsorption is favored by higher pressure and lower temperature, the opposite process, the desorption is carried out at low pressure and elevated temperature. With desorption, the bond between the adsorbate and adsorbent is broken. Thus the adsorbate is released and the adsorbent regenerated. Depending on the type of adsorbent/adsorbate system and the adsorption operating conditions desorption occurs in different ways that may also be combined [4.51, 4.521 :

Steam

Active carbon -

Heat storage Support grid

Distillate

Fig. 4-19. Discontinous fixed bed adsorber types, using active carbon. Representation according to WIRTH[0.1, Vol. 21.

4 Adsorption

312 0)

Purified gas

Purified gas

A

R

R a w g as

b)

Purified g a s

Purified gas Adsorbent

Downcorner

nor row-mesh sieve t r o y

t Adsorbent Raw gas

out

Fig. 4-20. Adsorber types. Representation according to BRAUER[4.10, 4.521. a) Discontinously operated adsorber with pellet-type adsorbent in packed beds of large volume, random packed in a thin ring layer or flat, thick layer b) Adsorber with multistage “well-mixed” fluidized bed

4.6 Desorption, Regeneration of Loaded Adsorbent C)

313

“PI u g - f Io w” f Iui d iz a ti on t r a y s . plan view

Purified g a s Adsorbent in

Straight path

Downcomer Narrow reversing path

fluidization t r o y

Raw g a s

d)

Adsorbent

Purified

Raw 9 0 5 -

i

ei

Raw

Purified g a s

Fig. 4-20. (continued) c) Continously operated adsorber with multistage (plug-flow-like) fluidized bed. 1 , 0 Input and output of adsorbent. d) Continously operated fluidized-bed adsorber with straight path. e) Continously operated adsorber with cross flow moving bed.

314

4 Adsorption

Temperature Swing Process Here the temperature of the loaded adsorbent is increased indirectly using a heating surface or by direct contact with hot gas thereby reducing the loading capacity of the adsorbate. The temperature swing process is often applied to molecular sieves where water, hydrogen sulfide, carbon dioxide, etc. are the adsorbates and with active carbon or silica gel, loaded with hydrocarbons. Pressure Swing Process Once the loading phase is finished the raised pressure in the adsorber is reduced by pressure release or evacuation, which then decreases the absorbability of the adsorbate on the adsorbent. The released adsorbate may be stripped using a strip

,-Paraffins Kerosine

gas or exhaust. The pressure swing process is faster than the temperature swing desorption process. For example, this procedure is applied to remove steam and carbon dioxide from an inert gas with molecular sieves or silica gel as an adsorbent, or according to Fig. 4-1 to remove nitrogen from air.

Displacement Desorption Here, an additional desorption medium, which is preferably adsorbed by the adsorbent displaces the adsorbed component. The desorption medium has to be chosen in such a way, that it is easily separated from the adsorbate. Figure 4-21 shows an adsorption plant for the separation of a mixture of normal and iso-

n-Paraffins

Fig. 4-21. Separation of n-paraffins from a kerosine cut by adsorption. Representation according to RUHL [4.49]. A Adsorber (loading of a 50 nm molecular sieve by n-paraffines from a kerosine fraction) c1 Column (separation of i-paraffins as bottom product) B Displacement component receiver, components are n-pentane or n-heptane D Desorber (displacement of n-paraffins by displacement component) c 2 Column (separation of n-paraffin mixture into long-chain paraffins as bottom product and displacement component as the top product) Condenser C EV Evaporator H E Heat exchanger ~~

References

paraffins. n-pentane and n-hexane serve as the displacement agent. 0

Extraction (Desorption into a liquid phase) The adsorbed adsorbate is extracted from the adsorbent surface by means of a solvent and is removed as a solution. Since the solubility of the adsorption substance in the solvent normally increases with increasing temperature, a temperature increase for extraction leads to an increase in the desorption rate. The extraction method is particularly favored by adsorption systems with weak bonds between the adsorbent and adsorbed component, provided the solute absorbs well, and if a temperature swing process can not be used due to thermal sensitivity of the adsorbate. A n example is the extraction of adsorbed sulfur using carbon disulfide as the solvent.

References [4.1] WEDLER, G.: Adsorption. Eine Einfiihrung in die Physisorption und Chemiesorption. Verlag Chemie, Weinheim 1970. S. R. : Ad[4.2] HAUFFE,K., and MORRISON, sorption. Eine Einfuhrung in die Probleme der Adsorption. w. de Gruyter, Berlin 1974. [4.3] MANTELL, C. L. : Adsorption. McGrawHill Book Comp., New York 1951. [4.4] BRATZLER,K.: Adsorption von Gasen und Dampfen. Steinkopff, Dresden 1944. [4.5] GREGG,S. J., KING, K. S. W.: Adsorption. Surface Area and Porosity: Academic Press. [4.6] WIRTH, H. : ,,Absorption," Ullmanns Encyklopadie der technischen Chemie, 4th Ed., Vol. 2. Verlag Chemie, Weinheim 1972.

315

[4.7] JUNTGEN,H.: VDZBer. 253 (1976). [4.8] WEYL,R. : Grundlugen der Adsorptionstechnik. Unpublished script. [4.9] KAST,W.: Chem. Zng. Tech. 53 (1981) 3, 160- 172. [4.10] BRAUER,H. : Staubjournal 100 (1983) 6, 39-51. [4.1 I] MERSMANN, A.; M~NSTERMANN, U., and SCHADL, J.: Chem. Zng. Tech. 55 (1983) 6, 446-458. [4.12] RICHTER, E., and KNOBLAUCH,K.: Chem. Tech. (Heidelberg) 13 (1984) 9, 11-15. [4.13] KRILL,H., and WIRTH,H.: Chem. Ing. Tech. 46 (1974) 18, 757-762. [4.14] Company report: ,,Abluftreinigung durch Adsorption," T 1320/4.79, Lurgi GmbH. [4.15] Company report: ,,Kontisorbon-Verfahren," T 1477/6.84, Lurgi GmbH. [4.16] BRAUER,H. W.: Chem. Tech. 13 (1984) 7, 18-23. [4.17] Fa. UHDE:Anwendung yon Aktivkohle in der Abwasserreinigung. Printed matter La V15 19 2000 81. [4.18] BOHL,F.: Umwelt Tech. 81, 6-8, 80. [4.19] EISENACKER,K., and HOSENFELD,E.: ,,Abwasserreinigung mittels Aktivkohle, " 9. Arbeitstagung Verfahrenstechnik,Graz 1980. [4.20] CAPELLE, A., and DE VOOYS,E (eds): Activated Carbon, A Fascinating Material. Norit N.V., Amersfoort 1983. [4.21] Company report: ,,Silica Gel," Fa. Grace. [4.22] Company report: ,,KC-Trockenperlen," Fa. Kali-Chemie. [4.23] Company report: ,,Baylith," AC 14600 Fa. Bayer Leverkusen 1981. [4.24] MENGEL, M. Chem. Tech. (Heidelberg) 10 (1981) 11, 1135-1139. [4.25] MERSMANN, A. : Thermische Verfahrenstechnik. Springer-Verlag Berlin, Heidelberg 1980. [4.26] KOGLIN,B., LESCHONSKI, K., and WULF, A.: Chem. Zng. Tech. 47 (1975) 1, 21-24. [4.27] WIRTH, H.: Staub Reinhalt. Luft 36 (1976) 7, 288-292. 14.281 COLLINS,J. J. : Chem. Eng. Prog. Symp. Ser. 63 (1967) 74, 31. [4.29] KRILL,H. : Staub Reinhalt. Luft 36 (1976) 7, 298-302.

316

4 Adsorption

[4.46] SCHNEIDER, H. : Dissertation, TU Berlin [4.30] ROSEN,J. B.: Ind. Eng. Chem. Process 1978. Des. Dev. 46 (1954) 8, 1590-1594. [4.47] SCHMLDT, B. : Dissertation, TU Berlin 1978. [4.31] JOKISCH,F. : Dissertation, T H Darmstadt [4.48] JUNTGEN, H. : ,,Physikalisch-chemische 1975. und verfahrenstechnische Grundlagen von [4.32] KAST,W., and JOKISCH,F.: Chem. Ing. Adsorptionsverfahren," HDTVortragsverTech. 44 (1972) 556-563. oyfentlichungen 404 (1978) 5-24. [4.33] DENGLER, W., and KRUCKELS, W. : Chem. [4.49] RUHL,E.: Chem. Ing. Tech. 43 (1971) 15. Ing. Tech. 42 (1970) 1258-1266. 870-876. [4.34] DENGLER, W., and BLENKE,H.: Erfah[4.50] RICHTER,E., KNOBLAUCH,K., and JUNTrenstechnik 8 (1974) 8, 239-245. GEN,H.: Chem. Zng. Tech. 50 (1978) 8, [4.35] JUNEEN, H., HARDER,B., and KNOB600-61 1. LAUCH,K.: Chem. Ind. 35 (1983) 1, K.-D., KLEIN,J., and JUNEEN, [4.51] HENNING, 38-41, 87-90. H. : VDI Forschungsh. 615 (1983). [4.36] MICHAELS,A. S. : Ind. Eng. Chem. 44 [4.52] BRAUER,H.: Chem. Ing. Tech. 57 (1985) (1952), 1922-1930. 8, 650-663. [4.37] ERGUN,S.: Chem. Eng. Prog. 48 (1952), [4.53] DE ACETIS,J., and ' ~ O D O S ,G. : Ind. Eng. 289-294. Chem. 52 (1960) 1003. [4.38] UHLMANN,H. J., and KRUCKELS, W.: [4.54] ,,Adsorption an Kohlefaser fur Abluft Verfahrenstechnik9 (1975) 1. mit organischer Belastung" Wasser, Luft [4.39] RUTHVEN,D. M., GARG, D. R., and Betr. 30 (1986) 5 , 40-42. CRAWFORD, R. M.: Chem. Eng. Sci. 30 [4.55] ,,Abluftreinigung mit neuer Kohlefaser(1975) 803-810. Adsorptionstechnik," Oberflaeche JOT [4.40] KRISCHER, O., and KAST, W.: Die 26 (1986) 9, 30-32. wissenschaftlichen Grundlagen der Trocknungstechnik,Vol. 1. Springer-Ver- [4.56] CANS,W.: Fortschr. Ber. VDI 3 (1986) 116, 1-131. lag, Berlin, Heidelberg 1978. A. : "Adsorption," Ullmann's [4.41] BRAUER,H. : Grundlagen der Einphasen- [4.57] MERSMANN, Encyclopedia of Industrial Chemistry, und Mehrphasenstromungen. Verlag SauerVol. B3. VCH Verlagsgesellschaft, Weinlander Aarau 1971. heim 1988. [4.42] MOLERUS,0.: Fluid-Feststoff-Stromungen. Springer-Verlag Berlin, Heidelberg [4.58] COULSON,J. M., RICHARDSON,J. F., BACKHURST, J. R., and HARKER,J. H.: 1982. "Adsorption," Chemical Engineering, [4.43] KRISHNAIA,K., and VARMA,Y. B. G. : Vol. 2. Pergamon Press, Elmsford, OxCan. L Chem. Eng. 60 (1982) 346-352. ford 1991. [4.44] BRAUER,H., MUHLE,J., and SCHMIDT, M.: Chem. Ing. Tech. 42 (1970), 494-502. H., and ASBECK, H. : Verfahrens[4.45] BRAUER, technik 6 (1972) 230-238.

Thermal Separation Processes: Principles and Design Klaus Sattler, Hans Jacob Feindt copyright 0VCH

Verlagsgesellschaft m h H , 1995

5 Drying

5.1 Concepts, Processes and

Examples Drying is the separation of a liquid from a moist product. The separation may be carried out in a rough mechanical manner without phase change using filtration, centrifugation, or pressing out. In particular, moisture is physiochemically separated from gas mixtures using special hygroscopic drying substances. In this case, the drying substance adsorbs the moisture, as in gas drying with a molecular sieve; or a chemical reaction takes place between the moisture and drying substance. For example, phosphoric pentoxide (the drying substance) is converted by water to phosphoric acid. Moisture may also become part of the water of hydration of the drying substance if it is free of water of crystallization and amorphous. Therefore, the substance behaves very hygroscopically while in contact with the moisture of the moist product. Moisture may also be removed from a moist product by supply of heat. This thermal drying is discussed as a thermal separation process in the following sections. Thermal drying consists of two steps. Firstly, heat is supplied by conduction, convection, or radiation from outside the product, or is generated inside the product. This heat is used to evaporate or vaporize the moisture out of the product. Secondly, the product phase and the steam are separated as follows. Steam is removed from the product, and if necessary, condensed outside of the dryer.

In thermal drying processes, heat and mass are simultaneously transferred. If the product surface is sufficiently wet, the drying rate is only a function of the mass and heat transfer at the surface (surface evaporation). If the moisture content in the product is below a critical moisture content, the drying rate is largely controlled by heat transfer and the movement of the moisture inside the product to the surface (drying inside the product). The hygroscopic behavior, pore structure, and thermal conductivity of the product are then decisive. Mechanical predrying is usually installed upstream of thermal drying, a typical drying stage consisting of the process units shown in Fig. 5-1 and a combination of these arrangements shown in Fig. 5-2. If a moist product with an initial moisture content X& (given in kg moisture per kg dry solid), is predried by mechanical means such as centrifugation or filtration to a moisture content XeI, the required cost increases with decreasing Xel. However, the energy input required in the subsequent thermal drying increases with increasing X e l . Therefore, the intermediate product moisture content, Xel, has to be adjusted to minimize the total expenditure for the mechanical and thermal drying. Thermal drying is carried out through convection, contact and radiation drying, and also for special cases such as vacuum drying, freeze drying (sublimation drying), and high frequency drying. Sometimes different process variations and dryer types are combined, depending on the drying behavior of the product and moisture content, to reach an optimized minimum cost.

318

5 Drying W e t feed

Exhaust gas

t

I

i

Fig. 5-1. Convection drying. -- Circulating convection drying

-

Suction cell

I

-

-

Table filter

fl"er%

*

0

-+!iI t

Push- type centrifuge

If I

--

Peel-type centrifuge

Decanter centrifuge

Plate feeder

k

Pneumatic dryer

-

I

A

t Fluidized b e d dryer

Disk dryer

'

p\

Vacuum plate

Meed t t

Vacuum disk dryer

319

5.1 Concepts, Processes and Examples

Figure 5-3 clearly illustrates the simplified principles of thermal drying. The heat (2 supplied to the product by conduction, convection, or radiation is used to heat the product and to vaporize or evaporate the moisture m,. Evaporative or convection drying are mainly used; contact or radiation drying are applied in principal with vacuum drying. Since with contact or radiation drying compared to convection drying, the necessary effort for environment protection (emission or immission level protec-

tion) and sometimes the energy input is lower, these variations become more important for practical applications. Thermal drying is an important thermal separation technique used in most technical fields, for example, drying of fuel, dyes, foodstuffs and luxury items. Due to the different properties associated with different moist products, it follows that there are many process variations and dryer designs. Mathematical modeling is made very difficult or impossible due to the different and

b)

Wet feed Rotary

--- -

1

Shaping

69

Spray dryer

Drum dryer

Belt dryer

Fig. 5-2. Common combinations of drying devices. Representation according to VORHOLZ[5.9]. a) Drying of free flowing solids b) Drying of slurries

I

Grooved drum drier

Disk dryer

Pneumatic dryer

320

5 Drying

-

2 -

Hot g a s is u s e d a s a heating medium a n d to carry away ihe vaporized liquid

5.2 Characteristics of the Moist Product, Movement of Moisture

WM

Vaporized liquid is withd r a w n under vacuum or b y using purge g a s

cl ////////~///////////////. 0

moisture content of 2-5 070. The final moisture of 0.2% is achieved in a fluidized bed dryer.

J EA

Vaporized liquid withd r a w n under vacuum or by purge g a s

WM

I'ig. 5-3. Principles of thermal drying. a) Convection drying b) Contact drying c) Radiation drying WM Wet material HS Heated bearing surfacc EA Radiation emitting area m,, Mass o f moisture removed from pi.oduct Q Heat

changing states of the moist product during drying. To date, no uniform design concept for dryers exists. Figure 5-4 shows a two stage polyvinyl chloride drying unit. The polyvinyl chloride, which is formed by polymerization, is mechanically drained, using a centrifuge and a pneumatic dryer, from an initial water content of 20-40% predried to a remaining

The product to be dried is characterized (Table 5-1) as

Solid: - free flowing (powdery, crystalline, fibrous, granulate, flaky, lumps) - pieces - flat Pulp or Paste: - high viscosity although able to be pumped (liquid of high viscosity, high viscosity gelatin, muddy) - high viscosity, not able to be pumped (thick paste, moldable pastes) Liquid: - solution - colloidal solution - suspension - slurry The product contains moisture (usually water) mainly in the form of adsorbed liquid, capillary liquid, liquid causing the solid to swell or as bound liquid. Adsorbed liquid is a thin liquid film spread on the outer product surface. The vapor pressure above the liquid film corresponds to its saturation pressure at every temperature. Capillary liquid wets the inner pore surface of a porous solid. During the drying process, the liquid has to be transferred to the product surface by capillary forces. With products having macro capillaries of size m, the vapor pressure still approximately cor-

321

5.2 Characteristics of the Moist Product, Movement of Moisture

I

sfI I I

I

I1 I

‘ -

J&

4-

1 -

P&j5”

I

2-

I

I I

I\

I I

/

I I

I

-12

-+L

1 Centrifuge 2 Pneumatic dryer with I water-cooled iacket. first drying stage Dry product 3 Air heater 4 Pneumatic dryer 5 Cyclone 6 Scrubber 7 Circulating tank 8 Pump 9 Product receiver 10 Fluid bed dryer 1 1 Blower Fig. 5-4. Two-stage polyvinyl chloride drying unit. 12 Cyclone Representation according to Babcock-BSH AG, Krefeld.

+

t

//

Classification of wet material handled and correspondend to commercial dryers'). ze term of ze ranges

mm Fine molecular

term in dustry (clay)

10-5

10.~

Colloid (macro molecular)

A) . liquid (colloidal solution)

10-3

10-2

10.'

1

Microscopic

1

I

(A) 1. liquid (molecular solution)

10

Macroscopic

Fine sand

1

Coarse sand

I I I

1

Fine gravel

1

I

Soot

I

Kaolin

Clay

!. gelantinous

A) . liquid !. viscous

(B) (Aerosol) fine dust pulverized crumbly

c

Paint

I

gment

Flue ash

I. gelantinous

B) Aerosol) 'ine dust iulverized rumbly

Coal dust

A) I. viscous L. sludgelike !. muddy 1. pasty 5. clay-like

C

Yeast

c

B) Aerosol) 'ine dust mlverized irittle

I

)

Corn starch

viscous 2. sludge like 3. muddy 4. pasty (doughy) 5 . clay-like 6. crumbly 7. brittle

B) Aerosol) 'ine dust iulverized :rumbly irittle umpy

rm

Wheat flour :A)

I.sludgy 2. pasty (doughy) 3. crumbly 1. brittle 5. pulverized

(B) (Aerosol) fine dust pulverized criimhlv

I

Molding sand

I -

Wood chips Wheat grair

1. mushy

1

corn

3

2. crumbly Pasta

ommercial continous roughput

m

1

1. Spray dryer Interlayer drum dryer

3

10-2

I . Spray dryei 1. 2. Interlayer drum dryer 2. Belt dryer with specia pretreatment unit

10-1

1

Spray dryer Interlayer drum 1.

i

4.

]

with special pretreatment unit

B) lust dverized :rainy irittle :rumbly umPY

]

Interlayer drum dryer Belt dryei

-

(A) 1. brittle 2. grainy 3. flaky 4. chip-like 5. fibrous 6. shaped 3) :rit xlverized xittle xumbly

A) . crum !. pellet shape objec 1.

(B) brittle grainy flaky chip-like fibrous shaped

1o3

1

Interlayer 1. drum 1 dryer 4. Belt dryer with special

system Air con-

Tricklerotary kiln dryer

I.

1. Spray drye 1. Spray dryer 2. Interlayer Interlayer drum drye drum dryer Belt dryer 2.1 Belt dryer

1.:

1

\

6.) 2.

5.

I04

Troughflow drye type Fluidized. bed dryer Air con-

Ove dry typ Thr flo dry typ

system Air conveying dryer Tricklerotary kiln dryer

Tricklerotary kiln dryer

tion according to KROLL [5.8] and Babcock-BSH AG, Krefeld. jects and endless films in the macroscopic range different kinds of baffle radiation dryers are used. mmon delivery form; (B) Mostly obtained final form. Common delivery form” and “usual commercial dryers with continous material throughput” the same numbers correspond to each other.

324

5

Drying

responds to the saturation pressure. With products having microscopic pores < 10-7m, the vapor pressure is lower than the saturation pressure. During the drying process this decreases because the moisture is bonded in the micro capillaries. The solid is termed “hygroscopic”. It takes moisture from the surroundings as long as the moisture vapor pressure is equal to the saturation pressure. Sweliing liquid not only wets the surface of the product but also causes the solid to swell and becomes part of the structure of the product. Moisture is generally colloidally bonded to the product. Removal of the moisture during drying causes the particle to shrink. Moisture may also be present as water of crystallization. Removing the water of crystallization requires exceeding a crystal decomposition temperature which is specific to the crystal and which is generally not obtained during thermal drying. If a hygroscopic wet product comes into contact with a vapor mixture containing steam, moisture is adsorbed until the adsorption equilibrium is reached. If, however, the wet product is in contact with a dry gas, the product releases steam until the moisture partial pressure in the gas corresponds to the vapor pressure inside the product. Adsorption and desorption equilibria are described by sorption isotherms, its courses depend on the type of product, the moisture content, and the moisture bond with the product (see Chapter 1.4.4.2). The course and time of the drying process depend on the drying conditions, on the temperature and moisture profiles developed during the drying process, and above all, on the moisture movement in the product. Moisture movement is governed by the properties, form and dimension of the product, and the type of moisture bond in the product. Some relationships are listed in Table 5-2 to describe different mechanisms of moisture movement.

5.3 Properties of Wet Gases, h-X Diagram In vacuum drying, the released wet vapor is exhausted. In convection drying and other processes under normal or over-pressure, the moisture is carried away by an auxiliary gas, which is inert with respect to the solid. The gas flows over the product or penetrates through the product. Additionally, in convection drying, the gas is used as a heating medium whereby heat is transferred by convection to the wet product. The evaporated moisture contents in the wet auxiliary gas, the absolute moisture content or the moisture load X of the gas is (5-8)

where mDis the mass of the moisture in the wet gas and mL is the mass of the dry gas. Including

the absolute moisture content, X is (5-10) Introducing the so called relative humidity p, (5-11) and therefore, (5-12)

325

5.3 Properties of Wet Gases, h-X Diagram

Table 5-2. Moisture movement in solid(s) with capillary size pores. Mechanisms of moisture movement. Representation according to [0.4,5.11. Moisture movement by capillary force and surface force If a capillary of radius r is dipped into a wetting liquid, the liquid rises in the capillary by height z above the surrounding liquid level.

Surface tension Liquid density g Gravitational acceleration The smaller the capillary the larger is z. Therefore, liquid is aspirated by narrow product capillaries from wider pores and transfered to the evaporation plane or product surface. The liquid flux riz,i transported by capillary force is, according to KRISCHER CT

el

K

Moisture coefficient as a function of wet product moisture content X , (determined experimentally [5.1]) A Product sectional area, perpendicular to the liquid movement en; Dry product density AX, Change of product moisture load along the transported distance As Vapor diffusion in pores According to KRISCHERvapor diffusion in vapor filled pores is

b ,u Ap 0

Motion coefficient, depending on the mechanism of vapor movement in the capillaries Resistance coefficient [5.1] Effective pressure or partial pressure gradient of wet vapor along the diffusion distance As

Wet vapor molecular movement in the pores (r Q A )

(5-4) r A

4

MD 0

Pore radius Mean free path length of a vapor molecule Mean equivalent pore diameter Molar wet vapor weight

Vapor diffusion by STEFAN (r % A )

b = -D. '-M D '

P

P-PD,m D Vapor diffusion coefficient in dry gas p Total pressure pD., Mean partial wet vapor pressure in the pores

(5-5)

(continued next page)

326

5

Drying

Table 5-2. (continued)

p D Partial moisture vapor pressure in the gas at the product surface po,O Wet vapor saturation pressure

1

0

Laminar and turbulent flow in the pores only filled with wet vapor @ + A ) b=-

eD qD

ca' . QD 32

'

VD

at laminar flow

(5-7)

For seldom-appearing turbulent flow see [5.1] Wet vapor density Wet vapor dynamic viscosity

In Eqs. (5-9)-(5-12), the nomenclature used is p D , p L partial pressure of the wet vapor and dry gas, respectively, in the wet gas atmosphere MD,ML molar masses of the moisture and dry gas P total pressure Po, D saturation pressure of the moisture at the same reference temperature as p D @ D , @,,D partial density of the evaporated moisture as undersaturated and saturated wet gas

The enthalpy h of (1 + X ) kg wet gas is composed of the enthalpy of the dry gas h , and the fraction of the evaporated moisture in the wet gas

if the enthalpy of the dry gas and the moisture in the liquid state are zero at O"C, and are referred to a temperature 19. Ah/,, is the vaporization enthalpy of the moisture, related to 0 "C. cP:&and c , , ~are the mean specific heat capacities of the dry gas and evaporated vapor, respectively, between 0 "C and 19.

If the wet gas is saturated with moisture, the partial pressure p D of the evaporated moisture is the saturation pressure p o , D . The saturation load X , follows from Eq. (5-12) with v, = 1. If the wet gas is oversaturated, the surplus moisture is distributed and carried in the wet gas in form of liquid droplets or small ice crystals ( X > X,), depending on the temperature. In the case of surplus liquid in the liquid state the enthalpy of the oversaturated wet gas is calculated (liquid mist range) as

For the case of surplus moisture in the solid state (ice mist range) the enthalpy of the oversaturated wet gas is

where cP,w,c , , ~ mean specific heat of liquid and frozen liquid between 0°C and 19 Ah,, melting enthalpy of the moisture

5.3 Properties of Wet Gases, h-X Diagram

Changes of state of the wet gas are conveniently demonstrated in MOLLIER’S h-X diagram [5 .lo- 5.121. On this diagram, the enthalpy of the wet gas h is presented as a function of the absolute moisture content X with the parameters being temperature B and relative humidity p. In order to obtain a sufficiently large range for the important areas of undersaturated or overheated wet gas, an askew coordinate system is chosen. The X axis is so inclined, that the 0°C isotherm of the wet, unsaturated gas is horizontal. The structure of the h-X diagram is explained in Fig. 5-5. The enthalpies h for chosen moisture load X and temperature L9 are calculated by Eq. (5-13). State points of ihe wet gas with equal enthalpy h form isenthalpic lines in the diagram with h = const. of slope i.e., inclined downward to the right. With B = const. and variable X in the diagram, the isotherms are straight lines which slope toward the isenthalpic lines 0

c

w9

tI

’ cp.0

.9

h

h I

327

lCP.L.3

\h

~ V+AhLg , ~ in the subsaturated area

cp, w *

8

c ~ , 19-Ah,,, ~ .

in the liquid mist area in the ice mist area

If p = 1 is substituted into Eq. (5-12) with different temperatures L9, and the related saturation pressure curve determined, a line of equal relative humidity p = 1 is obtained in the h-X diagram. This line subdivides the state areas of homogeneous, subsaturated, overheated wet gas and the state area of the heterogeneous, oversaturated wet gas, the “mist area”. On the saturation line, with p = 1, the isotherms change their inclination. Additionally, lines of equal relative humidity, SO called p lines, are also calculated by Eq. (5-12). The h-X diagram additionally contains a border scale, in which A h / A Xare related to a pole P. This is used for changes of state of the wet gas for both enthalpy h and the moisture load X.

X---t

I

Fig. 5-5. Mollier h, X-diagram. hi Isenthalpic lines 8, Isotherms q, p-lines of equal humidity

Adiabatic saturation lines, f l K lines, also of constant temperature in the diagram, are discussed later. The h-X diagram is valid for a definite pressure. Conversion to other total pressures is discussed in [5.10]. In Table 5-3 some practical examples and useful illustrations are presented on how to use the h-X diagram. Figure 5-6 shows a h-X diagram for wet air under a total pressure p = 1 bar.

328

5 Drying

Table 5-3. Sample applications for the MOLLIERh,X-diagram. Example, illustration in h,X-diagram

Comment

1. Representation of wet gas state

In the h,X-diagram the wet gas state point is clearly described by at least two of the following properties:

2. Dew point determination of wet gas

If a wet gas of constant vapor contents (mass fraction) is cooled the relative humidity (o increases until it reaches the value 1 at temperature d T . d T is called dew point. Moisture condenses by further cooling below the dew point. Fog is formed 0 Initial wet gas state before cooling 0 State point of just-saturated wet gas

XI

(5-16)

p,, is the corresponding saturation pressure to d,,

X-

3. Heating and cooling of wet gas

m Mass flow rate of wet gas mL Mass flow rate of dry gas Required heat flux Q,2 for cooling or heating. Q12

= mL.. fh, - h2)

(5-17)

Condensed moisture during cooling into the wet bulb area (point 0 )

Am, = I+ZL . (Xi - X,)

x-

CO Cooling HE Heating

c

(5-18)

5.3 Properties of Wet Gases, h-X Diagram

329

Table 5-3. (continued)

Example, illustration in h,X-diagram

Comment

4. Mixing of wet gas flows

If two wet gas flows m , and m2 are mixed adiabatically the following balance equation results: Overall balance m, + m2 = f i Drying gas balance mL,l+ r i ~ =~mL , ~ Moisture balance mL,,. XI + mL,2 X 2= mL . X , Heat balance mL,, . h, + mL,2. h2 = mL h,*

-

(5-19) (5-20)

(5-21) (5-22)

The mixture point M may be determined by the aid of X , and h,.

(5-23) (5-24) In the h,X-diagram M is easily determined by using the Lever rule. Which is:

(5-25) After the distance I between the state points 1 and 2 is withdrawn from the h,X-diagram, now I, may be calculated. 5 . Addition of moisture to wet gas

MS Mixing stage

Because the state points of wet gas liquid or vaporous moisture are infinite in the h,X-diagram, they may not be entered. Addition of moisture can not be treated as in Example 4. The course of the mixing line has to be determined in a different manner. Moisture balance mL,,. XI + mF = mL,,* X , (5-26) Heat balance mL,, h, + mF*h, = m , , * h*, (5-27) Solving both equations for mF or m F . h, and division of the second equation by the first equation gives (5-28)

The humidity enthalpy hF is marked on the side scale (point Q , for vaporous moisture and Q2 for liquid moisture, qualitatively). The line Pol P and Q1 or Q2, respectively, gives the direction of the mixing line. Mixing points M, and M2 are found at the intesection of the parallels PO1or PO2 through 1 with the lines X , parallel to the ordinate according to Eq. (5-26).

*

In the heat balance the enthalpy h has to be referred to the mass of drying air.

. 5-6. Mollier h,X-diagram for wet air. Representation according to KNEULE[5.2]. Isotherms 8 p-Lines

~

____

Isenthalpic lines h Adiabatic saturation curves 8/,(wet bulb temperature)

5.4 Mass and Heat Transfer in Convection Drying

331

Table 5-4. Required physical characteristics of air and water to evaluate the enthalpy equations. Specific heat cp at p = 1 bar and different temperatures of air ( c ~ ,and ~ ) water ( c ~ , ~ )

- 50

0 50

100

150 200 250 0

1.0055 1.0056 1.008 1.012 1.018 1.026 1.035

1.855 1.858 1.864 1.872 1.881 1.892 1.905

Additional data Air: Molar weight ML = 29 kg/kmol Water: Molar weight MD= 18 kg/kmol Evaporation enthalpy 0 "C Melting enthalpy Mean specific heat of liquid water Mean specific heat of ice

Table 5-4 gives necessary substance properties (physical characteristics) for the important aidwater drying system, required to evaluate Eqs. (5-10)-(5-15). The MOLLIER h-X diagram is not the only diagram in practical use. Details of other diagrams are found in [0.17] and [5.13]. With the knowledge of the appropriate substance properties for several moisture systems with dry gas h-X diagrams are calculated and plotted as ready-to-use programs [5.14].

5.4 Mass and Heat Transfer in Convection Drying In convection or evaporative drying, heat is transferred from the preheated drying gas to the wet product. This heat is used to evaporate moisture from the product into

Ah,,, = 2500 kJ/kg Ahs,/ = 333 kJ/kg cp, = 4.19 kJ/(kg . K) cPE = 2.05 kJ/(kg K)

the gas phase. The drying gas picks up moisture and cools down during the evaporation. The product loses an equivalent amount of moisture and may warm up slightly. If a drying gas flows over a sufficiently wet product with capillary pores and if the partial pressure p D of the moisture is lower than the saturated pressure po,D of the moisture inside the product, moisture evaporates from the product surface into the drying gas (surface evaporation, evaporation in the lSt drying stage, see Chapter 5.5). Since the amount of moisture evaporated is delivered from the inside of the particles to the surface by capillary forces, the product surface behaves like a free liquid surface with respect to evaporation. Therefore, the evaporation rate ~ D i.e., , the evaporation rate of moisture from the product surface, depends only on the mass transfer coefficient /3 and the difference between the wet gas composition on the product sur-

332

5 Drying

face x , and in the gas bulk x,. It is also a function of the type of moisture and the gas condition. Therefore,

molar mass of the moisture (kg/kmol) A product surface (m2) x0,xL molar fraction of the evaporated moisture directly on the product surface and in the gas bulk (for the derivation of Eq. (5-29), see 15.71). MD

or in simplified form For adiabatic surface evaporation, the drying gas delivers the required heat flow Q, which depends only on the heat transfer coefficient a , the driving temperature difference f l L - fl,, and the properties (specific heat and heat of vaporization) of the moisture :

(see also Fig. 5-7) where

p Q

mass transfer coefficient (m/h) =

mean molar density of the wet gas (kmol/m3) total pressure (bar) mean absolute temperature of gas boundary layer (K) Y

'

p T,

Tm

(5-31) b)

a)

Wet vapor flow

Heat flow

6

Drying gas

f

Q

mD

1

4

Drying gas

I

xL

u

Q,

C 0

c

VI .-'D

c

.-+-0

C

0

0 .c _

c L

0

0

Q

1

ul

L

0

Q Ill

C

L

+I

C

2

/ / / / / / / /'

Temperature 8 Moisture concentration x

tc

Fig. 5-7. Saturated and unsaturated surface drying, temperature and humidity concentration curves. a) Surface evaporation, 1'' drying period b) Unsaturated surface drying, drying of inside product (2nd and 3'd drying period) CF Capillary flow VD Vapor diffusion

5.4 Mass and Heat Transfer in Convection Drying

333

it follows from the exponent in Eq. (5-34)

or in simplified form

a

where heat transfer coefficient (W/ (m2 * K)) product surface area (m2) specific heat (kJ/(kg K)) and

which has the mean value of 1.3 for the air/ water system.

vaporization enthalpy of the moisture (kJ/kg) temperature of the drying gas 80,dL directly on the surface of the product and in the gas bulk (K or "C) (for the derivation of Eq. (5-31), see [5.7]).

A a = - - - - - - . thermal conductivity of the @ "P drying gas (m2/h) density (kg/m3), specific heat e, Cp, A (kJ/(kg - K)) and heat of conduction (W/(m - K)) diffusion coefficient (m2/h)

The adiabatic product surface temperature, or equilibrium temperature, do may be calculated by Eqs. (5-29) and (5-31). The heat flow Q required to evaporate the moisture flow riz, at the product surface has to be considered and if the drying gas state (L~L,xL)is known, then

The solution of Eq. (5-34) and therefore the determination of the adiabatic equilibrium temperature do may be calculated by a graphical iteration method. Lines of equal adiabatic equilibrium, or product surface temperature L9, are liquid mist isotherms extended to the undersaturation or overheated areas in the h-X diagram, for Le = 1 and small moisture loads, as derived from Eq. (5-34), [5.7]. It also follows from a simple inspection of an element of the product surface area completely covered with moisture (Fig. 5-8).

a A cp,D, Ahl,g

Q = riz,'

-

Ah,,g

(5-33)

and

(5-34)

In Eqs. (5-35) and (5-36)

~~

mL. X , h

I

WG

mL,X+dX.h+dh

L 7

with

(5-35) WP

and the Lewis number Le

a Le = D

Fig. 5-8. Balance scheme and wet gas state func-

tion. (5-36) WP Wet solid product WG Wet gas

334

5 Drying

If wet gas flows over the surface element, the steam content X is changed by an amount dX due to the absorption of moisture. Heat is supplied to the wet product and used to evaporate the moisture mL dX and to superheat to a temperature dL - dL9. The enthalpy of the gas is changed by dh. A heat balance then gives +

= UiZL *

riz,

.

(h + dh)

+

dm,

(5-38)

= Ij2L

. ( X + u)

(5-39)

Division of the appropriate modified balznce equations gives the change of state of the wet gas during the evaporation Process (5-40) Upon integration a linear relationship is obtained, h = cp,w .8,. X

t

h

+ const.

(5-41)

The change of state of the wet gas is, as shown, controlled by the surface temperature of the wet product Lp,. During the evaporation process the wet gas state points lie on a straight line in the h-X diagram, of slope cP,+,.. Lp,, an extension of a mist isotherm into the undersaturated area (Fig. 5-9). With adiabatic convection drying, the temperature at the product surface is not absolutely constant. It increases slightly while the wet gas temperature continually decreases. Corresponding states of the wet gas in the bulk and the product surface are described by the state points Li and Gi which lie on a straight line of slope according to the previous deriv-

L,,

. . ., I,,,

. . . L,, Wet gas state points in the

gas kernel G,, . . . G, . . . G , Wet gas state at the product surface LK Wet gas &ange of Slate line r9, Equilibriunl or wet bulb temperature

ative. After an infinite period of exchange, the wet gas and product reach an equal temperature, the equilibrium temperature or wet-bulb temperature 8,. The true change of state of the wet gas during evaporation is presented by a slightly curved line from L , to L , = G,, however, in drying practice the bend of the curve may be neglected. During the drying process the temperature L9” at the product surface is regarded as being approximately constant and is set equal to the wet-bulb temperature L9,. Therefore, the change of state of the wet gas during the evaporation process is approximately described by a line of constant wet-bulb temperature. 3

cP,w.8,

- x + const.

(5-42)

If the evaporative drying at the product surface is not adiabatic, for example, additional heat is supplied by radiation, this has to be considered in Eq. (5-38). The change of state of the wet gas is influenced by the

5.5 Drying Kinetics, Course of Drying, Drying Time

335

efficients a and p but also by the “driving forces” dL - 6, and X u - X L , and the properties of the wet gas and moisture. Additionally, liquid conduction, moisture diffusion to the drying front, moisture steam diffusion from the drying front to the surface, and hygroscopic moisture bonds to the product, have to be considered. A mathematical description of the mass and heat transfer during the drying process from inside the product can only be formed with certain simplified assumptions. In general, the drying course and drying rate have to be found experimentally. Fig. 5-10. Wet gas state in convection drying in

an h,X-diagram. LW Line of constant wet-bulb temperature 8, ED Change of state during convection drying including heat supply

additional heat flow, by an amount X , - X and by the mass transfer coefficient p on the right side of Eq. (5-42). This is expressed as a line having a smaller inclined course presented in the h-X diagram (Fig. 5-10).

5.5 Drying Kinetics, Course of Drying, Drying Time

The kinetics of drying are the change in the mean moisture content of a product, and the mean product temperature during the drying time. Therefore, kinetics describe the time course of the drying process. This drying course is influenced by the kind of wet The nomenclature in Eqs. (5-38)-(5-42) is product and the moisture bond, by the choh, X enthalpy and steam load (moisture sen drying process, and the operating conload) of the wet gas dition during drying. c , , ~ specific heat of the liquid moisture The course of drying is characterized by at the product surface plotting the moisture content of the proddm, evaporated amount of moisture in uct as a moisture load X , (given in kg the observed surface element moisture per kg dry product) against the The intensity of the capillary liquid con- time t (Fig. 5-11 a). Furthermore, the curve duction in a product with capillary pores is clearly defines the drying rate dX,/dt, i. e., only sufficient with moisture loads of the the change in the moisture content of the product X , > XG,K(see Chapter 5 . 9 , and product as a function of time (Fig. 5-11 b). if moisture is delivered to the surface at the To determine the required time tg for drysame rate as the evaporation rate to main- ing, it is convenient to plot the drying rate tain a certain moisture content at the sur- as a function of the moisture load of the face. If the moisture load falls below XG,k, product. This is the actual drying rate curve the evaporation zones in the second and (Fig. 5-11 c). Knowledge of the drying rate is essential third drying stage move inside the product for design and selection of the dryer. Dur(Fig. 5-7). The mass and heat transfer are ing the drying, heat is transferred to the wet now not only controlled by the transfer co-

336

5 Drying

xG.k1

product. This is used to vaporize the moisture bound in a liquid state to the product, and also to overcome the bonding energy. With no heat losses during the drying process, the heat carried away by the evaporated moisture is equal to the heat supplied to the process. This “equilibrium” of supplied and withdrawn heat at the drying front of the product controls the course of the drying. (The drying front is the level in the product up to which point the product capillaries contain moisture in a liquid state. Above the drying front steam diffuses in the direction of the product surface). In a short warm up period, the wet product with initial moisture content X G, , is heated to the operating temperature. A negligible amount of moisture has already been evaporated (Fig. 5-11, area AB). In the socalled first drying stage, area BC, the moisture content decreases linearly; the drying rate is constant. The first drying stage lasts until the quantity of evaporated moisture is different to that delivered to the product surface by capillary moisture conduction. Point C, the critical moisture content, representing the end of the first drying stage, with load XG,kof the product, is found experimentally. The higher the drying rate and the thicker the product layer, the higher the load, XG,k.X G, kalso increases with increasing mean diameter of the product capillaries. In the first drying stage, the thin film of liquid covering the surface, and some of the bound moisture in the macropores, evaporates. The drying rate only depends on the condition of the drying gas and the mass and heat transfer coefficients. For the case of convection drying, the amount of evaporated moisture lizD in unit time is (see Chapter 5.4)

w:7 ---__

XG,ti XG.E

\

F

C)

Fig. 5-1 1. Drying rate curve of porous solids. a) Product moisture content X , (dry basis) vs. time t b) Drying rate as a function of time c) Drying rate as a function of product moisture content Nonhygroscopic material ---_ Hygroscopic material

5.5 Drying Kinetics, Course of Drying, Drying Time

where pD , o and pD,L are the partial pressures of the evaporated moisture in the drying process, directly at the product surface and in the gas bulk, respectively. If with reducing moisture content X,, the rate at which moisture is delivered to the surface by capillary condition is less than the evaporation rate, the drying rate decreases, from point C . However, at C the second drying stage begins. The drying front now moves away from the surface to the inside of the product. The drying rate is now controlled by the heat conduction at the dry product surface, by the capillary moisture conduction from inside the product to the drying front, by moisture diffusion from the drying front to the product surface, and by transfer of the evaporated moisture into the drying gas. With increasing distance of the drying front from the product surface s and therefore with time t the drying rate decreases, since mass and heat transfer with increasing s are hindered by additional resistances s/Ddr and s/&, in the second drying stage ( D d r , Adr are the coefficients of diffusion of evaporated moisture in the dry product above the drying front and thermal conductivity of the dry product, respectively; R D is the individual gas constant of the moisture). Therefore, in a unit time the evaporated moisture rate Yiz, is m,=-.

1 RD.T

PD,~-PD,L

1

,U*S

P

D

-+-

.A

(5-44)

where ,u is a resistance factor depending on the thermal conductivity Ad,., and D is an effective diffusion coefficient, found experimentally [5.1]. With nonhygroscopic products, the second drying stage ends with complete drying of the product ( X , = 0, point F). Hygroscopic products can only be dried completely if the drying gas contains no mois-

337

ture. In practice this is not the case. Therefore, for hygroscopic products a third drying stage is included (area DE, dashed line). This starts at point D, where maximum possible hygroscopic moisture content XG,H is reached at every point in the product. (A product is hygroscopic if the bound moisture has a lower vapor pressure than the pure liquid moisture at the same temperature). XG,His the moisture content of the product in sorption equilibrium with the surrounding moisture saturated drying gas. The drying rate in the third drying stage decreases further to zero, at which point the equilibrium moisture content XG,E is reached. XG,Eis the residual moisture content in sorption equilibrium with the surrounding drying gas. Sorption isotherms, found experimentally (see Chapter 1.4.4.2 and Fig. 5-12), describe the lowest possible remaining moisture content of a product with a given drying gas state. Wet products without a pore system, such as gels, dough, and paste show no pronounced inversion points on their drying rate curves. During the drying process, the rate of drying decreases continuously from an initial value to zero, when the residual moisture content is reached. For individual cases, it is possible to calculate drying rate curves [5.2], although it is favored to find the course of the drying experimentally under operating conditions in a pilot scale dryer. Experimental arrangements, for example, are described in [5.I 5 - 5.171. Figure 5-13 demonstrates the influence on the drying rate of granulate shape, kind of air flow, air temperature and velocity, on a belt dryer using granulated particles (clay, pigments, plastic filler, etc.). In order to design a dryer, the required drying time tg to dry a product, with an initial moisture content X,,,, to the final content X G , 2 , must be known. The required drying time tg is obtained from the drying rate curves (Fig. 5-11 c). Therefore,

338

5 Drying

I

2

I a=C o n s t a n t

I

I

z E

0

02

I

I

I

//I I

I

I

OL 06 08 10 IRelotive v a p o r p r e s s u r e ratio (re1 humidity)y

~

*

---

---c t _ _ _ -

Fig. 5-12. Equilibrium moisture content

loyer adsorption ILonymuiri

the drying rate go over a time interval dt with the moisture mass rn,, dXG evaporated from the product of surface area A is

-

go

=

rnG, T *

dx,

A . dt

(5-45)

If go can be found as a function of the moisture content of the product X,, integration for the drying time tg with go and the boundary conditions XG,, and X G , 2 gives

and rnG,1 the mass of the wet product. Integration of Eq. (5-46) can be carried out by numerical approximation methods or by graphical means, by plotting l/g, against X , (Fig. 5-14). Linearization of the drying rate curve, gD(XG)for the first drying stage (area X G> XG,kl)corresponds to go

=

c,

(5-48)

and in the second drying stage (area X,,,, < X G < xG,kl) is (5-49)

where mG,Tis the mass of the dry product

m ~T =,

rnG, 1

1 + XG,l

(5-47)

With fixed product moisture contents, X,,,, and X G , E and , slope k, direct integration gives a simple approximation of the drying time tg

339

5.5 Drying Kinetics, Course of Drying, Drying Time a)

I

12mmQ

60 -

135'C

h

8mmQ

25mm

50-

10[kg/lm2~h)l

lO0OC

4%

800C

30xl0x6mm

30-

&-t 11

20-

0 2 1 6 810 WG

[%I

-

20

0 2 1 6 810

2L

W G [%]

-

20

24

. .

i

2 m/sec 1,5rn/sec

-

11

c

1

//I

2 .

0 0 2 1 6 8 1 0WG ["/a]

-

20

21

l0I

0

0 2 1 6 810 Wt

[%]

-

20

1

21

Fig. 5-13. Dependency of the drying rate on granulate shape type product, air flow, air temperature and air velocity in a belt dryer. Representation according to Babcock-BSH AG, Krefeld. Influence of the granulate shape on the drying rate (wL= 1.5 m/s, pL = 0.1, gL = 8 0 T , m,/A = 24 kg/m2) Influence on the drying rate of air flow passing through or over the bed (wL= 1.5 m/s, pL = 0.1, f l L = 100°C, m,/A = 12 kg/m2) Influence of the air temperature on the drying rate (wL= 1.5 m/s, pL = 0.05, mG/A = 24 kg/m2) Influence of the air velocity on the drying rate (pL = 0.025, rYL = lOO"C, m,/A = 24 kg/m2) Drying rate Product moisture content

340

5 Drying

t

1/g,

G ' ,E

XG.kl

'G.2

XG

-

G ' .1

Fig. 5-14. Determination of total drying time f,. X, Product moisture content iu Drying rate

[g=-.

%,T A Xc.2 xG - xG,E

I

(5-50)

and therefore

-*

In

(XG,l

-XG,kl)

xG,kl

-xG,E

x G , 2 - xG, E

I

1 f -

k

'

(5-51)

(Exact calculation of the drying time tg for various conditions and different drying behavior of selected wet products over the individual drying stages is given in [5.1, 5.21).

5.6 Convection Drying With convection drying, the required heat is transferred to the wet product by convection. The heating medium is preheated drying gas, generally air, but also containing inert gas or flue gas. To increase the drying effect, the drying gas not only flows over but also through the product, it fluidizes the product, and conveys the product pneumatically or flows around the disperse product phase.

5.6.1 Drying Gas and Heat Requirements in Convection Drying Consider a continuous adiabatic single stage dryer as shown in Fig. 5-15. The wet product flow, he,,is dried from an initial moisture content XG,,to a final state X C ;, ? .

5.6

Convection Drying

341

\ Fig. 5-15. Dry gas and heat consumption of

‘y!

dryers.

BL Blower GP Gas preheater

BL

DR Dryer

During drying, the amount of moisture removed is

where mG,= is the mass flow of the dry product and X , the moisture content of the product (kg moisture per kg dry product). Moisture is carried away by the drying gas. The mass flow rate of the drying gas mL does not change, but the total mass flow and the moisture load change

Following from Eqs. (5-52) and (5-53), the drying gas requirement rizL is

mL=

m D

x2-4

mG,l -

- mG,2

x,-x,

(5-54)

The specific drying gas requirement (“specific gas need”), related to the exchanged moisture I is

When drying using flue gas, as the heating medium, the total mass flow rate of all noncondensable components are summed to k L(e.g., nitrogen N,, nitric oxide NO,, oxygen 02,carbon dioxide CO,, carbon monoxide CO, sulfur dioxide SO,, etc.). If fuel of known composition is burned completely with an appropriate excess air coefficient,

c + 0, c02 2 H2 + 0 2 2 H2O +

+

0,

s + 0,

+

so,

n

-+

m C 0 2 + - H,O 2 (5-56)

the flue gas flow rate m and flue gas composition are calculated from the combustion stoichiometry.

where

mB

fuel mass flow rate m, mass flow rate of the combustion air mR, moisture flow rate of the wet flue gas

342

5 Drying

The moisture flow rate of the flue gas mRFis the sum of the water content of the fuel and combustion air, and the water generated during the combustion reaction. If flue gas cleaning stages are installed between the combustion and drying stages, their effect on the mass flow rate and composition of the flue entering the dryer has to be considered. With discontinuous operation of the convection drying process, the moisture content XG and moisture load X of the drying gas are time dependent. Considering the moisture mn(t) removed from the product and carried by the continuous flow drying gas, mD(t)= -mG, T .

dXG(t) ~

dt

,

*

= Yi?,

+ m,, 1 hG, 1 + Q = . h, + mG,2- hG,2+ Q v

.

by high-temperature oil

. Q = mo

*

From this, the heat required by the dryer, Q is found. (The energy requirement to release the moisture bound in hygroscopic products is disregarded in Q as are the heats of crystallization of solids out of the solution and the heat taken up by infiltrated air entering the dryer). The heat requirement decreases with decreasing gas requirement mL.The lower the exit gas temperature f12 and the higher the entry temperatures 8, and VG,, of the drying gas and product the lower the gas requirement. Further discussion of Eq. (5-59) can be found in [5.2]. The heat requirement Q transferred in an air or gas heater is provided by condensing steam

~* Aflo , ~

.

Q = m~ * Hu,B* qF

(5-61)

(5-62)

where in Eqs. (5-60)-(5-62)

&,I

(5-58)

(5-59)

c

or by combustion of fuel, mainly tion gas

(x2( t )-x,) h H D l

hi

(5-60)

= hHDAh,,,,

=mL.

where X 2 ( t )is the steam load of the drying gas leaving the dryer. A simplified heat balance over the dryer including the gas preheater gives with the nomenclature used in Fig. 5.15 yi2L

Q

c ~ , Afl, ~ ,

4,B VF

mB

mass flow rate of steam, high-temperature oil, and fuel specific heat and allowable cooling of the oil net calorific value of fuel combustion efficiency of the preheater

The time-dependent heat requirement Q(t) for a discontinuous convection dryer is given by

(5-63)

The heat capacity HG of the product is

also the moisture content X , and the enthalpy h, of the drying gas leaving the dryer are time dependent like the temperature. Differentiation of Eq. (5-64) with respect to time, gives

5.6 Convection Drying

(t) -

dHG

dt

- ~ G , T ‘

(c,,~,~+ X c (t) cp,w).

“‘“1

.-d8Gd t(0 +cp,W*8G(t)*p

dt

(5-65)

where c ~ , ~cp,w , ~ , specific heat of the dry prod-

343

If the wet product enters the dryer at the wet-bulb temperature 8K, with adiabatic convection drying the change of state of the drying gas follows the line of constant wetbulb temperature & (see Chapter 5.4). The drying process is demonstrated in the h-X diagram shown in Fig. 5-16. The specific heat requirement q is found directly from the diagram border scale Ah/AX for given entry and exit conditions of the drying gas.

uct and the moisture

Q ( t ) must be appropriately adjusted by controlling the heat supply to the air(gas) heater, to obtain a desired decrease in the moisture content XG(t) or course of the product temperature 8, (t). Further treatment of the discontinuous drying process is found in [5.19]. The specific heat requirement q of the convection dryer related to the exchanged moisture follows from Eq. (5-59) and by considering Eq. (5-55) 4=

P

h2 - hl 5 X,-&

where Q v is the heat loss of the drying unit. The positive sign for the second term on the right hand side of Eq. (5-66) accounts for cocurrent flow of the drying gas and product, the negative sign applies for countercurrent flow. Considering an “ideal dryer”, by neglecting the heat loss Q,,, and the difference between the latent heats of the entering and leaving product, it follows from Eq. (5-66) that 4=

t

h2 - hl

x,- Xl

(5-67)

5 - 6 2 Steps in Energy Saving Drying processes are one of the most energy-intensive process units. By considering the total economic efficiency of the drying unit during planning, appropriate energy saving and recovery steps have to be considered. Table 5-5 gives an overview of some of the possibilities for energy recovery/energy saving in thermal drying processes, particularly in convection drying. Figure 5-17 shows a schematic of a convection drying unit with waste gas heat recovery to heat the inlet gas. The change of

344

5 Drying

Table 5-5. Selected possibilities for energy recovery and energy saving in drying processes [5.1, 5.2, 5.20- 5.22, 5.241. 0

0

0

0 0 0

0 0

Prior to thermal drying product moisture content is reduced by mechanical dewatering (filter press, centrifuge, separator). The optimum moisture content with respect to economic efficiency results from total cost of predewatering and thermal drying. Reduction of heat losses by radiation, conduction and convection in the drying unit by the aid of improved heat insulation. Heat recovery by heat exchange between - exhaust gas and fresh gas (Fig. 5-17) - dry product and fresh gas (usually only low recovery efficiency) In steam-heated dryers use of vapor or steam, for example, hot water generation (water as product moisture and direct heat transfer). Use of vapor heat during drying of solvent wet products in an inert gas circulation. In a heat-pump cycle (Fig. 5-20) use of exhaust gas heat and vapor condensation heat [5.1061, In stage drying from the main dryer use of exhaust gas heat to heat up fresh gas in a predrying unit. In alternating load operation exhaust gas humidity is controlled by means of an exhaust gas flap valve [5.23]. In drying processes using flue gas as heating medium preheating of combustion air and, if necessary, product by exhaust gas. Adding of exhaust gas in the drying processes using flue gas (especially suitable to control the temperature of flue gas entering the dryer). Exhaust heat recovery by expansion and compression of exhaust gas (Fig. 5-19).

state of the drying gas is demonstrated in the h-X diagram in Fig. 5-18. Neglecting heat losses the heat requirement Q of the drying unit shown in Fig. 5-17 is

The heat flow Q, recovered in a heat exchanger is

By neglecting the heat contribution of the product, the heat recovery efficiency coefficient qR is (5-70)

and simplified FIR=--

82 - d3

&-@,

--& - $-,81 8R

(5-71)

(without considering the contribution of the moisture X to the wet gas enthalpy and the temperature dependency of the specific ) . is the total heat heats cp,Land c ~ , ~QH supplied to the drying gas corresponding to the temperature increase from d 1 to f l y . The optimum temperature 1.9, and, therefore, the most favorably achieved temperature @R and maximum heat recovery efficiency coefficient qR, result from minimizing the total costs of the heat recovery and gas heater [5.25]. Figure 5-19 shows a convection drying unit with waste heat recovery using expan-

5.6 Convection Drying

t

$1

FG

Fig. 5-17. Convection drying unit with waste gas heat recovery. DR Dryer G H Gas heater HR Heat recovery BL Blower FG Fresh gas EG Effluent gas H G Hot gas H A Heating agent W P Wet product D P Dry product 1, R, V, 2,3 State points in the h,X-diagram in Fig. 5-18

sion and compression. In a turbine, pressure-released gas is free of the moisture, which condenses during the expansion. After compression, the gas is ready for reuse. Assuming operation without heat losses, only mechanical work W is supplied (5-72)

345

Fig. 5-18. Wet gas state in convection drying with effluent gas heat recovery according to Fig. 5-17. 1 . ..R Fresh gas preheating by heat recovering R . . .V Fresh gas heating in the gas heater V . . . 2 Drying without heat losses 2 . . . 3 Effluent cooling by heat recovery

(W,, WEare used compression and gained expansion work, respectively; q v , vE are the efficiencies of the compressor and turbine, respectively). According to Fig. 5-20, a convection drying unit may also be combined with a heat pump cycle [5.27]. Waste gas leaving the dryer releases heat Q A , at a low temperature level to the vaporizer of the heat pump in which the gas loses moisture. In the condenser, the drying gas absorbs heat QH at a higher temperature level TH and flows back to the dryer. The rating E of the heat pump, including the compressor power Nv,is

c,

QH E = __

N V

(0.35 . . . 0.65).

TH TH - TA

(5-73)

5 Drying

346 a)

v MS

I

b'

I

1

h

I-

t

X-

SM

5-19. Exhaust gas heat recovery by expansion and Schematic Change of state in the h,X-diagram 1 . . . 5 . . .M Dryer Turbine M , .. 2 Compressor 2...3 MS Moisture separator 3 . . .4 WP Wet product 4...5 DP Dry product SM Separated moisture

Fig. a) b) DR TU CP

DR

4 W P

II

DP

5.6.3 Variations of Convection Drying If fresh air from the surroundings is used as drying gas, it is often necessary to mix some

compression [5.2, 5.261. Adding of little fresh gas to exhaust gas Exhaust gas expansion in the turbine Moisture condensation in the separator Compression of the drying gas Moisture pick-up during convection drying

Fig. 5-20. Exhaust gas heat recovery in a heat pump cycle. DR Dryer

of the used air, or waste gas, with it (Fig. 5-21). In this operating mode, called air circulation, or drying with fresh and recycled air, climatic variations in the state of the fresh air can be balanced by adjusting

5.6 Convection Drying

9

t

/

h

:1

P

the mixing ratio of fresh and recycled air. The specific air requirement I and specific heat requirement q remain unchanged with the same product moisture load m, compared with fresh air passed once through the dryer (Fig. 5-16). They only depend on the entry and exit state of the drying air. The exit temperature of the air from the heater f l v with recycled air, is lower than the temperature f l E when only fresh air passes through the dryer. Therefore, a more gentle thermal drying process is achieved. The compression work has to be considered

347

Fig. 5-21. Drying with fresh

and recycled air. Process scheme (a) and h,X-diagram @). BL Blower AH Air heater DR Dryer " FA Fresh air EA Exhaust air RA Recycled air -+- Single stage drying TIC Air ratio control by temperature

by determining the entrance temperature to the air heater f l A H

(VM, V A H ,P M , P A H , T M , TAH are the ~ 0 1 ~ metric flow, pressure and absolute temperature of the mixed air before and after compression, n is the polytropic exponent). The state point M of the mixture consisting of fresh and recycled air lies on the

348

5 Drying

n.

line The location of the point is fixed on this “mixing line” by the mixing ratio of the fresh air flux m,,, and the recycled air flux mL,uaccording to the Lever law (see Example 4 in Table 5-3): (5-75)

A4 is also found from a mass balance over

the mixing location. Since the wet product has to be gently thermally dried, the drying gas cannot exceed a maximum temperature g,,. A multistage drying operation is required (Fig. 5-22). The drying gas has to be heated to d, before entering the next drying stage. With stagewise drying, the required heat input q and needed drying air 1 are the

same as with single stage drying, with a single pass of air with the same moisture content of the product m, (dashed line in Fig. 5-22). Circulating drying and stage drying may also be combined in circulating stage drying. The advantage is that the state of the air can be adjusted to the respective product condition and high gas velocities allow good mass and heat transfer. Wet product and drying gas can flow in cocurrent flow, countercurrent flow, and cross flow through the dryer (Fig. 5-23). With cocurrent flow the entering hot gas makes contact with the wet crude product, thereby causing high initial heat and moisture transfer driving forces. The dry product leaving the dryer comes into contact with the cooled exhaust gas, which avoids damage of thermally sensitive products. With counter-

t - --

b)

F,-

--QE Ah lAX

t

h

P

Fig. 5-22, Two-stage drying unit, process scheme (a) and h,X-diagram (b). BL Blower GH Gas heater DR Dryer -+- Single stage drying

L

5.7 Contact Drying

349

No differences occur between cocurrent and countercurrent drying, as long as the product surface is at the wet-bulb temperature. With cross flow, the drying gas flows through, around, or carries away the wet product, which partly results in very intensive drying. Both the wet and dry product must be insensitive to high temperature. Cross flow drying is particularly useful if very short drying times are required.

5.7 Contact Drying

Fig. 5-23. Different types of convection drying, wet material and wet gas flow. a) Cocurrent drying b) Countercurrent drying c) Cross-current drying d) Multistage drying e) Muitistage drying with recycled gas _ _ _ Wet material Drying gas Gas heater ~

current flow, the entering hot gas contacts the dry product; therefore countercurrent drying of thermally sensitive products is not possible. The leaving waste gas flows over the wet fresh product, in which case it is possible that the gas temperature falls below the dew point and so condensation of moisture on to the product surface may occur. Under the same conditions, a lower final moisture content of the product is reached with countercurrent drying.

In contact drying heat is essentially transferred by conduction from heated walls, and if necessary, from the internals of the dryer, to the wet product. If the heat is exclusively transferred via the contact surface and there is no inert gas in the area accepting the moisture vapor, this is termed pure evaporation drying. Moisture evaporates at the drying front, and the moisture vapor is exhausted under vacuum, if necessary. Otherwise, if additional heat is transferred convectively from the hot gas to the wet product, a combined convection and contact drying takes place. Moisture evaporates at the drying front, diffuses through the dry product into the drying gas, and both are then withdrawn from the dryer. During contact drying the wet product rests on a heated base or is moved by stirring, or agitating with a shovel or circulation devices over the heated base. If the wet product is first moved in a thin layer over a heated surface, the product is preheated, and some moisture evaporates. The actual drying follows by evaporation of the remaining moisture, after which the product is superheated to the exit temperature. Figure 5-24 shows the moisture content X , and temperature t9 of the product

350

5 Drying

If the heat transfer coefficient kps is found by measurement or model calculations [5.28], the required preheating length zps may be calculated by Eq. (5-76) where wet product flux entering the dryer (kg/h) c ~ , ~specific , ~ heat of the wet product temperature at which the moisture 8, evaporates entry temperature of the wet 8, product U dryer circumference for heat exchange temperature of the heating medium dH (steam)

ti^^,^

QH

The heat transferred in the evaporation section QT (actual drying section) is ES

5s

Fig. 5-24. Wet product moisture (a) and temperature (b) as a function of heated surface length in contact drying. PS Preheating section ES Evaporating section SS Superheating section

Temperature Dryer length X , Product moisture content (9

z

plotted against the length z of the heated surface. In the preheating zone, the transferred heat Qps is

where mc,Tis the mass flow rate of the dry product, XG,ffand XG,ware the moisture contents at the inlet and outlet of the dryer, Ahl,, is the evaporation or bond enthalpy of the moisture, and k , is the heat transfer coefficient which is particularly dependent on the moisture content of the product. By means of this equation the actual required dryer length for drying zTcan be calculated. Eqs. (5-76) and (5-77) are generally valid for steam heated contact dryers with heat transfer from the internals of the dryer to the wet product. Only the individual design of the heat transfer areas z - U for each dryer must be considered. In contact drying, the heat transfer coefficient k , for a product piled up in front of a heating surface, is -kT =

(5-76)

1

-+-+-+aH

'W

(5-78) SF, 'TG

A,

5.8 Radiation Drying

The individual heat transfer mechanisms are : 0

0

From a heating medium to the heating surface (heat transfer coefficient a,) and then through the wall of the heating surface (wall thickness 6,, thermal conductivity A), Through the dry product layer (thickness of layer, ,a thermal conductivity An) and the wet product layer (thickness of layer AFG, thermal conductivity dFG) to the product surface.

kT is determined empirically under conditions as close as possible to the operating conditions. To fix the thermal conductivity of porous wet product or dry product, see [5.1]. To intensify the heat transfer in a contact dryer, the wet product is periodically agitated across the heating surfaces using an internal stirrer (e. g., rotary shaft with blades) or by rotating the complete dryer. The agitation frequency increases with increasing revolutions and with it the heat transfer coefficient. The heat transfer coefficient k , for the case of contact drying with mixed wet product on the heating surface [5.2], is

1

k, =

1

6,

-+--taH

(5-79)

1 aB

where aB is the limiting heat transfer coefficient between the product and the heating surface for short-term contact. It is proportional to the heat penetration number (A cp . and inversely proportional to the square root of the time between two product rearrangements. With incomplete mixing of the product t-''2 is replaced by t-", thus aB is

-

aB =

2 ~

fi

.v-

*t-"

(5-80)

351

(A, cp, e are the thermal conductivity, specific heat, and density of the product, respectively, IZ = 0.33 for a disc dryer with rotating blades, trough dryer; n = 0.22 for tube dryers, t is the time between two product rearrangements [5.29]). A dimensionless method is applied to the heat transfer coefficient as a function of rearrangement frequency and layer thickness of the product in a rotary tube dryer, disc dryer and tumble dryer for the case of preheated free-flowing product [5.30]. Contact drying is especially used if an overflow or through flow of the product to be dried by air or other gaseous heating medium 0

0

0

0

0

Is difficult or impossible (e.g., with doughy and pasty products) Influences the product quality, for example, thermal oxidation with convection drying of polymer chips with air as a heating medium Causes waste disposal problems, for example, treatment of toxic products Makes it difficult to recover the evaporated moisture, for example, with wet solvent products Causes safety problems, for example explosive, moisture vapor

Thermally sensitive products can only be treated by contact drying if low residence times of the contact surfaces are sufficient for the drying or if a heating medium at a low temperature is available.

5.8 Radiation Drying With radiation drying [5.1, 5.31, 5.321 some of the electromagnetic radiation energy emitted from the radiation source is absorbed by the wet product. The absorbed energy heats the product and evaporates

352

5 Drying

moisture. The heating not only takes place at the surface, but also inside the product. Therefore, fast drying without cracks forming on the surface, such as with ceramic products, and without forming a skin on varnished surfaces, is possible. The following are often used as radiation sources : 0

0

Briehl radiator ( A < 3 pm) e. g., tungsten wire lamp usually with parabolic formed reflector, infrared Briehl radiator made of quartz with a half cylinder highly polished aluminum reflector Dark body radiator ( A > 3 pm) e.g., electric dark body radiators made of thin steel tubes with internal electric heaters and reflectors made of aluminum or ceramics, catalytic low temperature radiators using a catalytic reaction of gaseous fuel below the ignition point as a radiation source, gas-heated metal walls in the temperature range 500-800 "C or ceramic plates with temperatures up to about 1100 "C (A refers here to wavelength)

q

2

=

CS

El CS

El E2

T , T, w

(5-82)

1 1 -+--1 E2

black body radiation constant (= 5.77 W/(m2 K4)) emission ratio of the radiator emission ratio of the wet product temperature of radiator and wet product incoming radiation number; considers the mutual arrangement of radiator and product surfaces

-

Since cl and c2 usually depend on the wavelength of the emitted radiation and the temperature, the emission spectrum of the radiator and Q must be adjusted to each respective product state during the drying process. This is done in each section of the dryer, through which the product passes. In radiation drying processes with a high rate of heat flow per unit area, more intensive and faster drying is achieved than in convection drying (rate of heat flow per unit area is ca. 16,000-40,000 kJ/(m2 h). Radiation drying is expensive because of The radiator material and temperature the high cost of electrical power or combushave to be chosen so that the maximum ration gas to heat the radiators. It is mainly diation intensity remains in the wavelength used for short drying times of thin product range of approximately 0.4-800 pm. Favorlayers (e.g., lacquer layer on metal sheets able ranges of absorptive capacity of the [5.33]), ceramic products with thin walls, wet product, depending on the moisture and thin paper or textile sheets. In the latter content, should also be covered. case, it is also used in combination with The radiation energy Q exchanged becontact drying, if necessary. (Dark textiles tween the radiator and the product is calcuwith a relatively high absorption capacity lated according to the STEFAN-BOLTZMANN can reach water evaporation rates of ca. law for parallel absorption and radiation 30 kg/(m2. h) and higher. surfaces of area A

5.9 Dielectric Drying where

C,,2 radiation exchange constant

With high frequency or dielectric drying [5.6, 5.34-5.371, the wet product forms a di-

5.9 Dielectric Drying

electric between the electrodes of a plate capacitor exposed to a high frequency electric field operated at a frequency of 2-100 MHz. Some of the field energy is absorbed to move ions in areas of low conductivity in the product and to direct dipoles. The movement causes intermolecular friction and with it, generation of heat inside the product. This is used to heat up the product and to evaporate the moisture. By controlling the high frequency voltage at the high frequency generator and by changing the air gap between the electrodes and product, the thermal power can be adjusted to the course of drying of the product. For example, to dry wood, the specific energy consumption is 1-1.5 (kW . h)/kg water. High frequency drying allows gentle thermal drying of the product. Substantial deformation of the product and shrinkage cracks are avoided. Therefore, dielectric drying is employed for gentle drying of high-grade products, such as fine wood, ceramic products, food products and luxury goods. For example, with rod electrodes arranged perpendicular to the flow direction, moisture gradients of paper in longitudinal and lateral directions are equalized. Therefore the paper quality is increased in the final drying stage after the main drying of the paper sheets [5.38]. The electric power P passing from a homogeneous electric field to a dielectric (which uniformly fills the field), is in the general case

353

(the wet product) unbound moisture evaporates off the accessible product surface A . With constant product temperature during the drying process, the drying rate go is P

where

U

voltage (V) active current (A) in phase with the voltage E field intensity (V/m) f frequency (Hz) C= so. A/s capacity of the plate capacitor (F) & dielectric constant of dielectric (Table 5-6) c0 = 8.854 . lo-'' (F/m) electrical field constant tan 6 dissipation factor (Table 5-6) 01 heat transfer coefficient (W/(m2. K)) 8,, tYL temperature of gas at the product surface and in the gas bulk (K or "C) 4,E evaporation enthalpy of the moisture (k J/kg ; k J/kmol)

I

Dielectric drying is usually operated in a frequency range of ca. 2-100 MHz. Otherwise, microwave drying is a particularly gentle and hygienic fast drying process for P = U . I = 2 . n . f.U 2 . C . & - t a n 6 (5-83) hydrated products in foods, luxury goods, and pharmaceuticals. Microwave dryers are and in the case of a plate capacitor of plate operated in the frequency range of 2450 k 50 MHz (see also Table 5-6). Amplified area A , and plate separation s is high frequency energy generated in highP = 2 n f . E 2 A, s E~ E tan 6 (5-84) vacuum electronic tubes (e. g., klystron, magnetron, amplitron) is delivered to the This energy is converted to kinetic energy in wet product via waveguides (waveguides the molecules of the dielectric, and there- with side slits or meander shaped, etc., fore into heat. If air flows over the dielectric, [5.11). +

-

-

-

354

5 Drying

Table 5-6. Principles and technical data of selected electrical dryers [5.1, 5.21. Radiation dryer (see Section 5.8) Energy for product heating generated by absorption of electromagnetic rays 0 0

Ultraviolet radiation dryer (High- and medium-temperature radiation dryer, Briehl radiator, 1 < 3 pm). Infrared radiation dryer (Low-temperature radiation dryer, dark body radiator, 1 > 3 wm).

Dielectric dryer (see Section 5.9) In a high frequency electric field the product to be dried forms the dielectric. 0

Alternating dielectric field drying Alternating dielectric field drying

Frequency (MHz)

Wavelength in vacuum (m)

Capacitor field (High frequency drying)

13.56 27.12 40.68

22 11 7.5

Dipole radiation field (Microwave drying) [5.107]

1000 2450 5725

Dielectric constant quencies.

0.3 0.12 0.025

and dissipation factor t a n 6 of different substances at different fre-

E

Substance *

E

10

100

3000

tan 6 10

100

3000 (MHz)

Air, vacuum Mahagony wood Paper China Water

1 2.17 2.86 5.82 78.20

1 2.07 2.77 5.80 78.00

1 1.88 2.70 5.51 76.70

0 0.31 0.51 0.115 0.046

0 0.32 0.66 0.135 0.050

0 0.25 0.56 0.155 1.570

Electron-beam dryer Wet product is exposed to an electron-beam generated and accelerated in a self-accelarated electron generator, i. e., paint drying or paint hardening [5.39]. 0

Technical data: Accelerating voltage Penetrating depth Required energy dose for hardening Suitable for the treatment of flat throughput.

ca. ca. ca. or

150.. .300 kV 120.. .300 pm 2 0 . . .500 kJ/kg film-type painted or coated products at large

Resistance dryer Wet product is the effective resistance in an electrical circuit.

* Reference temperature 25 "C.

5.10 Freeze Drying (Sublimation Drying)

5.10 Freeze Drying (Sublimation Drying) Freeze drying [5.40-5.441 is a vacuum sublimation drying process. At temperatures below 0°C and under vacuum, moisture sublimes from a frozen wet product directly from the solid to the gaseous state. Moisture is, therefore, withdrawn from the product. Firstly, the wet product has to be frozen from -15 to approximately -50°C. The freezing temperature depends on the wet product and the type of moisture; the temperature often has to be lower than the eutectic temperature (see Chapters 1.4.5.2 and 1.4.5.4). Deep-freezing temperatures normally cause a fine crystalline structure in the ice phase which is unfavorable for the drying process, compared with a coarse crystalline structure. The chosen freezing method depends on the type of product and on the product throughput. Products such as vegetables, fruits, etc., and crystalline products may freeze in a fluidized bed freezer with cooled air or in a freezer tunnel; liquid and paste-like products are frozen on drum or belt freezers operated with cooling brine or a cooling medium. Filling this product into bowls with subsequent freezing in a deep-freezing room is also possible (applies to batch operation or small charge sizes). The product freezes in direct contact with the cooling agent, in contact with a cooled base and/or by moisture evaporation under vacuum. The freezing rate and final temperature essentially influence the drying time and final quality of the product such as structure, consistency, color, and flavor. With a slow freezing rate close to the freezing point (freezing rate < ca. 0.5 cm/h) only a few ice crystal seeds are formed, which slowly grow to large crystals. This is favorable for the subsequent sublimation drying and to

355

maintain the flavor. With faster freezing rates, (> ca. 2-3 cm/h) fine grain crystals are formed. An optimized freezing rate is mainly determined experimentally as a function of the wet product, drying process, and operating conditions. It lies in the range from 0.5-3 cm/h for food products. During freezing most of the moisture is withdrawn from the product and the product is converted to a n icy state. The time required for freezing must be found experimentally; it may only be calculated with simplified assumptions. The frozen product is usually granulated, sieved and then charged to the dryer. Under vacuum it is then dried either discontinuously on heated plates or continuously while mixed and moved over a heated surface. The pressure for moisture sublimation is about 0.1-1 mbar and essentially depends on the moisture sublimation pressure curve; the partial pressure of inerts should not exceed 0.01 mbar. Sublimed moisture vapor desublimes to ice on a cooling agent operated condenser. Since there is no capillary moisture conduction to the product surface in the frozen product, no first drying stage exists. It starts with the second drying stage. The ice or sublimation front continuously moves inside the product. The required heat for sublimation is transferred to the product by conduction or radiation, or with a combination of both. Moisture vapor diffusion in the dried product follows the KNUDSENmolecular flow model (see Table 5-2). In the third drying stage residual liquid moisture desorbs to the inner and outer product surface from its bonded state. Since desorption occurs only after complete sublimation of ice, the heat transfer has to be reduced to avoid heating the product over the permissible limit. With sublimation, drying rates are low because the low allowable rate of heat flow per unit area (the heat transfer) controls the process. Since the required sublimation

356

5 Drying

time is proportional to the square of the layer thickness or the square of the particle diameter, respectively, the product has to be distributed in thin layers over the surface of the dryer. The actual drying time has been found to be similar to the freezing time by experiment. If all the moisture is in the ice state and the drying is carried out according to the course of the sublimation curve at a sublimation pressure ps and a sublimation temperature Vs, the sublimed vapor flux per unit area is (5-86)

where q is the allowable heat transfer per unit area and A h , , is the sublimation enthalpy. The required drying time fg of wet product (at rest) with the layer thickness 6 may estimated by (for the derivative see [5.401)

effective diffusion coefficient slope of the sublimation curve at dS

De

d T/dp

Frequently stirring the product with a shovel on a heated base produces a decisive decrease in the drying time, f,. Freeze or sublimation drying is carried out discontinuously in a chamber freezer or continuously in a vacuum disc dryer, vibrating film dryer, cascade dryer, or spray freezing dryer depending on the product properties [5.2, 5.40, 5.431. Figure 5-25 shows schematically the assembly of a chamber freeze drying unit. Figure 5-26 gives the flow sheet of a continuous 5-tunnel CQC (Continuous Quality Control) freeze drying unit. The high investment and operating costs of freeze drying are only worthwhile for high-grade, thermally sensitive products. Certain important properties of the products are kept such as flavor, taste, and color, and also certain ingredients such as proteins and vitamins. Sublimation drying is used to

where @I

X G . (I

$0

k A

density of ice initial moisture content of product porosity of the product total temperature gradient constant temperature of the heating medium temperature at the product surface overall heat transfer coefficient thermal conductivity of the product

L

Fig. 5-25. Chamber type freeze drying unit, design Leybold. Reprcscntation according to VOGELPOHLand SCHLUNUUR [&I, Vol. 21. 1 Wet material in containers on heated plates 2 Drying chamber 3 Condenser 4 Diffusion pump 5 Vacuum pump, first stage 6 Valves

5.11 Design of Dryers

357

Fig. 5-26. Continuous 5-tunnel CQC freeze drying unit by Leybold-Heraeus. Representation according to Leybold-Heraeus, Koln [5.44]. 1 Freezing chamber 10 Vacuum lock, exit 2 Charging chamber 11 Discharging station 3 Truck loading 12 Dry product outlet 4 Vacuum lock, entry 13 Scrubber 5 Condenser 14 Track system for truck transport 6 Pump equipment 15 CQC-hanging trucks 7 Gate (vacuum-sealed) 8 Rotary vacuum pump Lobe vacuum pump 8 Freczc drying tunnel 9 Plate heating system with lowered CQC-hanging trucks

0

dry and preserve certain food products and luxury goods, biochemicals, pharmaceuticals, and similar sensitive, high-grade products.

5.11 Design of Dryers From the discussion in Chapter 5.2, the existence of a number of different, technically important, wet products showing extremely different drying behavior is forwarded. An attempt to adjust the drying process to best fit the product behavior and throughput has led to a number of different dryer designs. In the following section, an overview of the classification of technical dryers, their selection, their operating mode, and design

is given. For further discussion, additional literature is cited.

5.11.1 Overview of Dryers, Dryer Selection and Design Criteria for dryer classification include: 0 Operating mode (continuous, discontin-

uous)

0

0

0

Operating conditions (surrounding pressure, vacuum and overpressure dryers) Appearance of the product to be dried (e. g., dryers for free-flowing, piece form, flat, liquid, paste, and dough products) Drying processes, means of energy supply (convection, contact, radiation, high frequency (dielectric), microwave, and freeze dryers, or dryers with combined energy usage, etc.)

358 0

0

5 Drying

Time required, drying time (short-term, medium-term and long-term dryers) Forming (shaping) of the product to be dried (granulating dryer, milling dryer, etc.)

In Table 5-1, frequently used wet products are related to important dryer designs. In Table 5-7, some technically important dryer designs are given, and classified by the drying process/motion of the wet product inside the dryer and/or motion of the dryer internals. The dryer designs presented are discussed thoroughly in [5.1 and 5.21. Typical application and design are pointed out in these references. Reference [5.45] gives in the form of a list, a “market overview” of dryers obtainable in Germany from 53 manufacturers, and 81 dryer designs to give the user a first insight. Information on dryer design, its primary field of application and some characteristic data enable an initial comparison. The solution of a drying problem starts with the selection of a drying process and dryer design. The basic information required is the behavior of the product to be dried during the drying process, the required drying time, and the first results from balancing (Fig. 5-27). Possible dryer designs can be found using a “solution catalogue”, which is set up as shown in Table 5-8. Common dryer designs are listed according to a variety of operating applications. The final selection of the dryer design is made after an economic comparison. Experiments in a pilot-scale plant under conditions as close as possible to operating conditions are often necessary for the selection. The dimensioning of the dryer is then quite simple if the required drying time tg (see Chapter 5 . 5 ) is known. The required dryer volume I/ is (5-88)

where I+Zc,l is the mass flow rate of wet product, ec,/is the density of the wet product, qF is the degree of filling, AQis the cross section and Z is the length or height of the dryer. With discontinuous operation, the times for charging and discharging, etc., have to be considered in addition to tg. The cross-sectional area A Q of the dryer is

This applies for the case of convection drying of wet product, with a transport velocity of we through the dryer and a drying gas (mass flow mL (1 + X I ) ,density eL)flowing over the product with an acceptable velocity. If the product is lifted up by the drying gas, for example with the airflow dryer, the allowable gas velocity w, is determined as the favored transport or conveying velocity by considering a tolerated dust carry over. (e. g., operational range of fluidized beds according to REH (Fig. 5-33), state airflow dryer). AQ is then

-

A Q=

h L (1 . +XI) WL * e L

(5-90)

A calculation for the first stage of a convection dryer is possible when the heat transfer coefficient a , and the volume specific product surface a as the phase boundary surface are known. In a differential dryer volume AQ . dz the gas temperature decreases in the direction of flow z by d19, with an approximately constant product surface temperature d o . While the gas moisture load increases, the corresponding moisture content of the product decreases (see Chapter 5.4). When neglecting the portion of the heat flow of the moisture load in the gas, the heat flow rate dQ transferred through AQ dz is

-

5.11 Design of Dryers

359

Table 5-7. Common commercial dryer designs. Convection dryer Product rests on rigid surface 0 Kiln dryer 0 Reciprocating dryer 0 Oven dryer, dry chamber with horizontal and cross flow of drying air 0 Jet dryer Product rests on movable surface 0 Canal dryer 0 Belt dryer (single and multibelt dryer, turbo-belt dryer, spiral-belt dryer, vertical belt dryer, loop and conveying belt dryers, short-loop dryer, tilting pan dryer, tower dryer) Product is moved by wipers 0 Disk dryer 0 Turbine dryer 0 Ring-tray turbo dryer Product moved by gravity 0 Drum-type dryer 0 Trickle dryer 0 Turbine shaft dryer Product moved by inertial force 0 Vibrating dryer 0 Chute dryer Product moved by drying medium 0 Float dryer 0 Fluidized-bed dryer 0 Conveying dryer 0 Convex dryer 0 Cyclone-type dryer with aggregate fluidized bed 0 Spiral-flow dryer 0 Spray dryer (for liquid product) Contact dryer Product rests on rigid surface 0 Cylinder dryer Product moved by stirring devices (wipers) 0 Plate dryer 0 Trough dryer 0 Screw-conveyor dryer 0 Blade dryer 0 Discotherm dryer 0 AP dryer (dough mixing dryer) 0 Druvatherm-plongh-blade dryer 0 Rovaktor (blade dryer) Product moved by gravity 0 Drum-type dryer 0 Forced-circulation dryer 0 Mill dryer Product moved by gravity and rotor 0 Drying tower 0 Horizontal moving bed dryer (continued next page)

360

5 Drying

Table 5-7 (continued)

Vacuum dryer 0 0 0 0 0

Vacuum drying chamber Contact vacuum dryer Tumble dryer, tilted-cylinder dryer Double-cone blender dryer Spray dryer

Radiation dryer 0

Infrared radiation belt dryer

High$requency dryer 0

High-frequency belt dryer

breeze dryer 0 0

0 0

Chamber dryer Vacuum spray dryer Tunnel dryer Cascade vibration dryer Plate dryer

Dryer with combined energy inpiit 0 0

0

Convection-contact dryer Convection-radiation dryer Contact-radiation dryer

-

- = ~ 9 8,) ~ - A y .dz . a =

dQ = a ( d L - 8,) dA =

a (

-

- yi2L . cP,L . d8L

(5-91)

On integration, this gives the required dryer length or height Z , for mass and heat transfer

Z = HTU, NTUG

(5-93)

As for mass transfer, the dryer length or height Z , is the product of the number NTU, and the height HTU, of transfer units. Eq. (5-92) may be used for drying in all drying stages if a and a are mean values over the total drying time tg, given empiri-

cally, and 8, as a constant temperature at the drying front. a also includes the specific heat flow inside the product. With constant 8, the temperature profile d L ( z ) of the drying gas in the dryer is

The total heat flow Q transferred along the dryer length or height Z , is

=

a A

Z . a - Aflm

(5-95)

5.11 Design of Dryers

Drying problem analysis Wet product throughput Batch size, batch frequency Wet product properties Type, form, consistency Initial moisture content Melting point. sticking point. softening point Decomposition and damage range Safety data Corrosivity Toxicity Desired properties 01 dry product Quality demand (final moisture content, purity, solubility etc.) Formulation features (powder. dusty or nondusty, agglomerate, granular, chips etc.)

361

7

If necessary, determined experimentally

-

Determination of characteristic data Process parameter, operaling mode Sorption isotherm Rate of drying curve, drying time Product properties during drying (physical characteristics. flowability, shrinking. particle size distribution, etc )

I

Criteria catalog Drying process Dryer design

If necessary, determination of data by experiment in bench scale Simulation of several process alternatives or dryer designs in laboratory scale

J.

Rough cost estimate Process selection Selection of dryer design

c

Dryer balance Mass balance Drying gas consumption Energy (heat) consumption Selection of steps to save drying gas and energy Disposal method optimization

If necessary, experiments in pilot scale at operating conditions as close as possible lo production conditions Basic design Scale-up, referred to the pilot scale Long-term operation experience Product sample for further treatment

I

Dryer dimension determination (Volume, cross-sectional area, length and heighl, contact area for heat exchange. installations. etc )

I

+

Design of peripheral equipment of the drying unit (Blower, gas heater, product charge and discharge devices. exhaust gas cleaning unit,

Fig. 5-27. Selection and design procedure of drying process and dryer design.

For the first drying stage, the local drying rate go is aexp

(-

a.AQ.z.a ,

m L . Cp,L

(5-96)

362

5 Drying

Table 5-8. Possible criteria catalog to select drying process and dryer design. Selection or evaluation criterion Wet product mode - free-flowing - pellet-like, flat - pasty - liquid (solution, suspension) - corrosive (chemical and abrasive) - solvent wet

Drying process or dryer design (examples) Convection dryer Contact dryer BD DD FB CD BLD PD FFD

+ + +

-

+

-

+ +* +

+

Operating form - continous - discontinuous

+

-

+ +

Drying time - short - medium - long

+ +

+ +

+

-

Gas flow - cocurrent - cross flow - countercurrent

+ + +

-

-

-

Operating cost

+

+

-

+ + +

Flexibility

+ +

-

+

-

Cleaning cost

-

+

-

Supervisory cost Maintenance cost

t-

-

+

+ +

+-

+

-

+ +

Space requirement - floor space - height Investment cost

Specific evaporation rate

-

-

Effect on product mode during drying process - keep mode - granulating - grinding

-

+

+ +

* With fluidization aid Abbreviations: BD Belt dryer; DD Drum dryer; FB Fluidized-bed dryer; CD Conveying dryer; BLD Blade dryer; PT Plate dryer; FFD Falling-film dryer + suitable/good/less + - limited suitability no or only very limited suitability/bad/high

5.11 Design of Dryers

Then the mean drying rate gD,, is gD,m =

a ~

Akld

*

(5-97)

A‘in

where in Eqs. (5-91)-(5-97), (see also Chapters 5.4 and 5.5) ‘ L , C I ~‘LL,o

mL

CP,L

Ah,,

temperature Of the drying gas at the entry and exit Of the dryer mass flow rate of the drying gas specific heat of the drying gas vaporization enthalpy of the moisture

The required heat exchange areas A , and A To f the preheating and drying sections in contact dryers are calculated by Eqs. (5-76) and (5-77). If the product state significantly changes during drying, the drying section has to be split into several stages to enable the mathematical treatment. The heat of dissipation by the power input is generally negligible. With a given mean drying rate gD,, (kg evaporated moisture per m2 heating area and h the length of the drying zone) of one drying section obtained from full-scale experience, A , is simply AT=

m G , T ‘ (xG, I

- *G,2)

363

5.11.2 Individual Presentation of Selected Dryer ’Qpes with Design Aids Very common dryer designs from Table 5-7 are briefly discussed with respect to their structure and operation. Hints for the design, characteristic key values and also further literature are included in individual descriptions.

5.11.2.1 Chamber Dryer The wet product in the form of lumpy solids, pieces, or a paste is spread out on wire gauze or sieve platedtrays, etc. A warmed drying gas flows over and passes through the stationary product surface. The gas is circulated by means of a blower, in which adjustable guide blades ensure an even gas distribution to each product layer. Chamber or tray dryers [0.6, 5.1, 5.2, 5.531 are operated discontinuously usually in gas circulation mode and are suited for small amounts of wet product. For drying, the required total drying area A corresponds to the total support area of the trays, which is

(5-98)

gD, in

where m G , T , XG,land x G , 2 are the mass flow, and the moisture content of the dry product at the entry and exit cross sections of the drying section of the dryer. In addition to this, methods for determining the dimensions of a dryer, a single dimension and, if necessary, the movable internals, require expert treatment of the apparatus-specific key values, which mainly result from operation experience. The “correct” dryer geometry and operating conditions follow from several design solutions with the demand of optimized economic operation of the complete unit.

(5-99) where

mg,/ mass of fresh product to be dried in

the operating time tb drying time, derived from the rateof-drying curve or found by experimentation, to reach the desired final moisture contents of the product. Possible times for charging and discharging, and the time for heating must be considered in tg mA mass of wet product spread per m2 tray area

tg

364

5 Drying

Chamber dryers are also commonly used as drying cabinets in the laboratory or on a pilot scale. Kiln dryers for lumpy, flat and solids pieces are also chamber dryers with huge chambers, from the point of view of process engineering. The product is spread on immovable or moving trays, or turned and moved over the trays using wipers or blades. 5.11.2.2 Tunnel Dryer

Tunnel dryers [5.1, 5.21 work with the same principle as the chamber dryer, but allow continuous operation to dry anything from large lumpy solids to pasty products. Trucks with trays on top of each other with evenly distributed wet product pass through a channel or tunnel, in which the drying gas flows over and/or through the wet product. 5.11.2.3 Belt Dryer

Anything from lumpy solids to pasty wet substances can be spread out in channels by means of a charger at one or more inlets, for example, with continuously moving endless belts made from cloth, steel plates or wire netting [5.1,5.2,5.47-5.491. Whilepassing through the fumigated drying channel, the product is dried in one or more levels (Fig. 5-28). The required belt area A is calculated as (5 -100)

where 6 is the thickness of the layer of the wet product on the belt and e is its mean density.

5.11.2.4 Multiple Plate Dryer

The free-flowing wet product passes through a drying tower with multiple vertically arranged trays where the hot gas flows over the product [5.1, 5.21. With plate dryers, the product is moved by means of revolving wiping blades, from the inside to the outside and vice versa and from plate to plate through the dryer (Fig. 5-29). Plate dryers are also operated as contact dryers under atmospheric pressure and vacuum [5.30, 5.461. This arrangement is used to dry free-flowing noncrusting products. The advantage is that the conveying and overturning or stirring conditions allow a narrow residence time distribution. With increasing revolutions of the main shaft, the revolving frequency increases and with it the drying rate. On the other hand, the mean product layer decreases and the mean residence time which is also important for the final drying, decreases. The relationship between product drying and conveying and several characteristic parameters are presented in an operating diagram [5.46]. In a ring-plate dryer the product is evenly spread out on annular trays by means of leveling blades, to obtain a layer of uniform thickness. The product remains on the tray for one revolution of the rotor before being moved to the next tray below by a wiper blade. The drying gas, circulated by externally mounted fans, flows over the product. In a ring-plate turbo-dryer, the product flows in a manner similar to the turbo drier through the drying sections. However, the drying gas is distributed over the product by centrally mounted turbine fans. 5.11.2.5 Rotary Dryer

Free-flowing wet product is charged to a rotating cylinder, which is installed at a small angle to the horizontal. The product is overturned and evenly distributed over the cross

5.11 Design of Dryers

365

a1

t

WM

EG

4

-

DM

I I

1

I -1. ~

4 1

HG

EG

I7

1

HG

section using internals (specially arranged lifting flights, see Fig. 5-30) [5.2, 5.18, 5.50, 5.511. The product then slowly moves through the dryer. The drying gas flows in the cylinder in countercurrent or cocurrent flow to the continuously moving solid. The heat required for the drying process is transferred by convection (e. g., the hot gas may be flue gas for materials such as limestone), or by conduction (heated wall of the rotating cylinder, electrical supply, or steam) and radiation through the wall and installations. If the revolving cylinder is also rearranged

Fig. 5-28. Three-deck band dryer. Design by Biittner, Schilde, Haas. Representation according to Babcock-BSH, Krefeld. a) Three-deck band dryer b) Controlled gas flow WM Wet material inlet DM Dry material discharge CB Conveyor belt F Fan HG Hot gas inlet EG Exhaust gas outlet

as a tube mill, the wet product is treated by a combination of milling and drying, a milldrying process. The required volume of the rotating cylinder VRis (5-101)

where rF is the degree of filling, hDis the moisture flow rate, and hD, is the evaporated moisture flow per m3 rotary cylinder volume.

366

5 Drying a)

WM

AI EG

b)

WM

I

:t I RS

GH

RS

PL

--c

FG t

.RD

Fig. 5-29. Plate dryer. Representation according to Krauss-Maffei AG, Munich. a) Plate dryer, schematic b) Wet material flow path WM Wet material charge (feed hopper) PL Plates RS Rotating stirring devices RD Rotor drive DM Dry material discharge FG Fresh gas EG Exhaust gas discharge G H Gas heater

The required length of the rotary cylinder may also be calculated with the NTU, HTU concept (see Chapter 5.11.1, [5.18, 5.521).

5.11.2.6 Fluidized Bed Dryer When drying gas flows through a solid bed of granular, or powdery, or short fibrous product, it behaves like a liquid (fluidized

ir DM

RD

1

state) if the gas velocity, related to the empty vessel cross section, reaches or exceeds a certain limit (fluidizing point, minimum fluidization velocity). Individual particles are suspended by the gas flow; the fluidized bed may be stirred like a liquid. The pressure drop of the drying gas while passing through the bed is approximately equal to the weight of the bed over the cross section (Fig. 5-31) [5.111, 5.1121. An increase in the gas velocity, above the minimum fluidization velocity, leads to an expansion of the fluidized bed and the formation of gas bubbles. With an increasing gas flow, the number and size of the bubbles increases. With very high gas velocities, individual bubbles are indistinguishable, and no defined bed surface is observed. This state is called a highly expanded or circulating fluidized bed. Due to the large carry over of solids, solids have to recharged to the bed by means of a cyclone. An equilibrium is reached between the solids carried over and those recharged, the

5.11 Design of Dryers

367

EG

WF

t

HG--

Fig. 5-30. Rotary dryer. Representation according to Babcock-BSH, Krefeld. WF Feed chute DG Driving gear GM Gear unit, gear mounting DS Rotary dryer shell RV Rotary valve DM Dry material discharge HG Hot gas inlet EG Exhaust gas discharge

Arrangements of lifting flights a) Radial flights for large, sticky or abrasive products b) Cruciform flights c) Cross-sectional flights for coarse product d) Quadrant flights for fine product

circulated bed is operating at steady state (Fig. 5-31). In a fluidized bed dryer [5.1, 5.2, 5.7, 5.66, 5.67, 5.1081, also called a fluid bed dryer, hot gas is evenly distributed over the cross-sectional area by means of a gas distributor (perforated plates, cap distributors, etc. [5.111]). Wet product is lifted over the entire cross section or fluidized. Due to the intensive contact between gas and product and good mixing behavior of the bed (“well-mixed”), the wet product is quickly dried (Fig. 5-32). Product movement in a fluidized bed is based on the quasi-hydraulic behavior of the system, i. e., with continuous operation, the level differences appearing between the feed and withdrawal zones cause product movement similar to a liquid flow. This kind of dryer is suitable for products which are readily fluidized (noncohesive) and free

368

5 Drying

I

a)

Ci

FB -

I

9

9

I1

f

Ape

AD

p I I

II

111

t a

' -*I

I

SB

j

I

FB

I

I

Fig. 5-31. Fluidized-bed ranges, pressure drop and heat transfer. Representation according to WERTHER[5.58] and KESTNER 15.721. .

A

I Fluidized-bed ranges a) Fixed bed (wg< w,,~) b) Minimum fluidization (ws= wnd) c) Bubbling (wn> w , ~ ~ ) d) Slugging (wg > wmfl) e) Fast fluidization (wg& w , ~ ~ ) SB Slumbed bed or fixed bed FB Fluidized bed CFB Circulated fluidized bed (pneumatic conveying)

I

i

PC

ICFB)

I1 Pressure drop versus superficial gas velocity MF Minimum fluidization; SCO significant carry over starts SB Slumbed bed FB Fluidized bed PC Turbulent bed, start of circulated fluidized bed (pneumatic conveying) NPD Narrow particle size distribution WPD Wide particle size distribution Ap Pressure drop 111 Heat transfer coefficient versus superficial

gas velocity wmf Minimum fluidization velocity w I Particle terminal velocity wg Superficial gas velocity a Heat transfer coefficient

5.11 Design of Dryers

flowing. Bed heights of 250-1,500 mm allow large residence times. Vibrating fluidized beds are mounted on plate springs which vibrate on a base frame with counterweights. Vibrators generate oscillations in the direction of product movement or conveying. Relatively small bed heights of up to 250 mm, and little backmixing of the fluidized bed (“plug-flow like”), in combination with the vibration, allow a narrow residence time distribution. Vibrating fluidized bed dryers are suitable for treatment of products which are difficult to fluidize or which tend to agglomerate or stick together. If additional heat exchanger equipment is put into the fluidized bed, heat transfer between the gas and heat exchanger and the product and heat exchanger, are superimposed to give the mass and heat transfer between the gas and the product. The transfer rate is significantly increased compared to a bed without a heat exchanger of the same area. The advantage is that the drying gas requirement is smaller, and hence the expenditure for exhaust gas treatment is reduced. The higher thermal efficiency also makes a higher evaporation rate at a lower temperature level favorable. Fluidized bed dryers with additional heat exchangers are operated with bed heights over ca. 400 mm, to treat fine granular products which are easy to fluidize. With dryers employed as spray-fluidized beds, the flowable product to be dried (solution or suspension) is sprayed onto an already dry product which acts as a receiver in the fluidized bed. During the following drying process, individual particles grow due to the sprayed product. Through a stirring effect and the convective mass transfer caused by the rising bubbles in a fluidized bed, intensive particle motion is observed. This is favorable for the mass and heat transfer between the gas and product and for the heat transfer between product and heat exchanger. Due to good solid mixing an almost uniform ternpera-

369

ture inside the whole bed is achieved. The large exchange area of gas and product particles is also favorable as is the liquid-like behavior that simplifies the handling of the product. The disadvantages are particle attrition, erosion of internals, and carry over of solids due to eruption of bubbles at the bed surface. Application of conventional fluidized bed technology with fines is problematical and limited economically, particularly with a specific light product having a high initial moisture content. The transformation of the fluidized bed principle into the centrifugal field offers some advantages [5.109]. Table 5-9 gives some indication of the fluid mechanics, the mass and heat transfer in fluidized beds, and their design. Fig. 5-33 shows a simplified operating range diagram according to REH I5.541-[5.56]. Gasisolid systems are classified into three operating ranges: fixed bed; pneumatic conveying; and an area which includes fluidization and classifying of solid particles. In Fig. 5-34 the operating ranges of a fluidized bed dryer and other dryers are shown. The Reynolds number Reg of the gas phase, calculated with particle diameter dp, is (5-102) and is plotted against a number including the Froude number, Fr

The parameter is the mean porosity of the bed E. A grid with lines of constant Archimedes number A r (5-104)

370

5 Drying

0)

Exhaust gas t o scrubber

bl

t I "

W e t product

"Plug- f I0w '* fluidized bed zone

Preliminary

Drying zone

t

t

Hot air

Hot air

t

Cold air

Dry Product

5.11 Design of Dryers

371

Wet feed Exhaust air t o cyclone

(1)

Side elevation

Adjustable weir

Hot air

I

Dried product

Plan view

(111) Plan view

Plan view

Feed

Feed

+ m t

Product

L7-

Product

Product

Feed

Fig. 5-32. Fluidized-bed dryer. and data of Babcock-BSH AG, Krefeld. Representation according to POERSCH Drying unit 1 Fluidized bed 2 Gas distributor 3 Charging device, wet material 4 Dust separator 5 Blower 6 Gas heater I Weir 8 Cooling section 9 Dry material discharge Fluidized-bed dryer Continuous “plug flow” fluid bed dryer: (I) straight path; (11) reversing path; (111) spiral path. Geldart [5.112].

5 Drying

372

Table 5-9. Indication of fluid mechanics, mass and heat transfer and design of fluidized beds.

Fluidized-bed cross-sectional area A , (5-106) Drying gas mass flow rate of density eE,kg/s m, Optimum fluidization superficial gas velocity, m/s wg Slumbed bed or packed bed height 2, of wet product Z , =mGLI .

t,=0.2

... 1 m

(5-107)

@s ’

mG,/

e, t,

Wet product mass flow rate, kg/h Wet product density, kg/m3 Required drying time, h

Gas phase pressure drop 0 Gas passing upwards through a polydisperse packed bed with minimum fluidization velocity wmf [5.57, 5.581

0

(5-108)

Gas passing through a fluidized bed Weight =

- particle buoyancy

Bed cross-sectional area

Gas passing through the gas distributor

a,, Emf

Zmf

wmf

ylg

-

A , * Zmf* ( 1 - emf). (e, - e,) g = const A, ( 5 -1 09)

ApD= (0.1 . . .0.2) . Ap, = 20 mbar (approximate value) Volume specific surface of the particles, m2/m3 Bed voidage at minimum fluidization (gas fraction, void fraction) Bed height at minimum fluidization velocity, m Minimum fluidization velocity, m/s kg * s Dynamic viscosity of gas, . m2

(5-110)

~

Minimum fluidization velocity wmf (Transition between slumbed bed and fluidized bed, see Fig. 5-31) [5.58]

(from Ap = Apn, Eqs. (5-108) and (5-109)) Kinematic viscosity of gas, m2/s.

vg Optimum fluidization velocity wg Determination by Fig. 5-34 based on operational experience or, for example [5.59] by W E= 7.5 . (5-112) at maximum heat transfer between fluidized bed and immersed transfer surfaces.

5.11 Design of Dryers

373

Table 5-9. (continued) Approximate value: wg = (1.5. , .2) . wmf for grainy, free-flowing particles wg = ( 5 . . .lo) wmf for particles with tendency to agglomerate

dp

(5-113) (5-114)

Mean particle diameter, m

Bed voidage emf at minimum fluidization 4 . m, (5-115) Emf = 1 x ' e, * d 2 . Zmj Mass of powder in a bed, kg m, d Bed diameter Height of incipient fluidized bed (determined experimentally or from expansion Zmf diagrams [5.60]) Fluidized-bed height Z , total height Zeffof fluidized bed and free board (disengaging space) (5-116) (5-117)

Zeff= z+ z , Reg, A r Reynolds and Archimedes number (see Eqs. (5-102) and (5-104)) & Bed voidage Free board height for allowable solid entrainment, m Z, Bubble formation, bubble growth, jet penetration [ 5 . 5 8 , 5.61, 5.111, 5.1121. Particle entrainment from fluidized beds [5.68, 5.69-5.71, 5.111, 5.1121 Mass and heat transfer in a fluidized bed 0 Estimate of mass and heat transfer coefficient in a fluidized bed without internals by empirical correlations (5-118) at Re, c; 100 [5.55] Nu = 0.3 . (5-119) at 0.1 < Reg< 15 L5.551 Sh = 0.374 . 0

(for a more detailed calculation see [5.59, 5.111, 5.1121). Estimate of the heat transfer coefficient a, between fluidized bed and immersed surfaces by empirical correlations

Nu, =

4

= 0.79

.Ar0.Z2.

( - y),.,, 1

-

[5.67]

(5-120)

10 < Ar < 2 . lo4 for For the case of horizontal, quadratic shifted tube bundles of tube diameter dR and of tube separation s (favorable values of a, are at wg = (2.5.. . 3 ) . wmf and s = (4.. .6) . dR),(for a detailed calculation see [5.59, 5.60, 5.63, 5.641). Dimensionless numbers Re,, Nu, Nu, and Ar are based on the particle diameter dp as the characteristic length Heat conductivity of gas, W/(m. K). I,

374

3

--Fr,

c

5 Drying

e,1 e, - e,

Re,

-

Fig. 5-33. State diagram of gaslsolid systems simplified for fluidized beds range of operation by REH [5.54-5.561. Representation according to POERSCH,Fa. Babcock-BSH AG, Krefeld.

5.11.2.7 Air-Flow Dryer, Pneumatic (Flash) Dryer [5.1, 5.21

and (5-105) helps to ease handling. (Nomenclature for Eqs. (5-102)-(5-109) is given in Table 5-9). With the knowledge of the particle size dp these diagrams provide an estimate of the fluidization velocity wg and the gas velocity for the other dryer types are found. Alternatively, with a chosen gas velocity wg the size dp of the particle is determined, which, for example, is either definitely carried over or just transported by pneumatic conveying. Figure 5-35 shows common residence time ranges for different convective dryers plotted against corresponding particle sizes.

In an air-flow dryer [5.7, 5.73, 5.741 powdery, granular, wet products are dried in parallel flow quickly and gently (Fig. 5-36) while being conveyed through an air-flow tube or ring channel [5.75, 5.761. The length of the dryer tube, the drying time and moisture vapor flow are decisive functions of the particle diameter of the product. Since the dryer has little product holdup during the drying, the operation may be adjusted for a rapidly changing moisture content and different feed flows. The advantage is also the low heat requirement and operational demand as well as the space saving vertical arrangement. The terminal velocity of the product particles is

375

5.11 Design of Dryers 10:

10

ps 1

3 . F r1 L

e,-e,

lo-’

lo-:

10‘:

lo-’

lo-:

-

10’

100

lo-’

10-2

Re,

102

lo3

1oL

Fig. 5-34. Ranges of operation of different convective dryers. Representation according to POERSCH, Babcock-BSH AG, Krefeld.

w,= f:.dp.-.-

e, - e g eg

g

c,

(5-121)

where c, is the resistance coefficient of a spherical particle, dependent on the gas Reynolds number Reg c w 3= q g + 1 y

(5 -122)

Mass and heat transfer in an air-flow dryer [5.74] are described using the Prandtl number Pr, and the Schmidt number Sc, (see Table 1-18), and are

Sh =

~

’’

dp = 2

D

+ 0.664 - SC;’/~Re;/2 (5-124)

With the relative velocity wre, between the gas velocity wg and the particle velocity w,, the Reynolds number is calculated

376

5

Drying

- 1 min

10-2

lo-’

1

d, [ m m l

-

10

Fig. 5-35. Ranges of residence time of different convective dryers. Representation according to POERSCH, Babcock-BSH AG, Krefeld. dp Particle diameter t Residence time

The nomenclature used in Eqs. (5-121)(5-125) is given in Table 5-9. The residence time tp of a particle in an air-flow dryer of length Z is (see also Fig. 5-35) (5-126) To dry a wet product with different particle sizes, the dryer has to act as a classifier. In a conventional air-flow dryer the classification effect is not sufficient: during the residence time in the first drying stage coarse

particles may not become dry while fine particles can reach too high a temperature. With other short-time dryers, the product motion is in the form of a vortex flow driven by the hot gas which increases the residence time. This can be achieved with hot gas jets tangentially entering the drying area (vortexflowdryers such as helical flow tube with nozzles and annular flow dryer [5.77]), with displacement internals including gas guide vanes (spiral tube pneumatic dryer [5.78, 5.79]), or with tapered inclined perforated screen (cyclone dryer [5.80]). The convex dryer [5.81, 5-82] combines the functions of drying and classifying. With a

5.11 Design of Dryers

t

I "

I -

.T

3e

C

b

' if

Fig. 5-36. Pneumatic dryer and helical flow dryer. Representation according to BabcockBSH AG, Krefeld. a Fresh gas fan b Gas heater c Helical flow tube with nozzles d Pneumatic tube e Exhaust gas fan f Circulating gas heater g Circulating gas fan h Dried material discharge by rotary valve i Wet material supply by feeding centrifuge

combination of a fluidized bed dryer and an air-flow dryer, i.e., a spinflash dryer, pastes, filter cakes, slurries, etc., can be processed [5.83].

5.11.2.8 Spray Dryer

In a spray dryer or suspendedparticle dryer (atomizing dryer) [5.1, 5.2, 5.7, 5.84-5.891

377

(Fig. 5-37), materials ranging from a liquid (solution, suspension) to a wet liquid-pasty product, can be sprayed as a liquid mist into the hot gas flow. In the case of a disc atomizer, the liquid product is pumped into a quickly rotating spray disc and leaves, driven by centrifugal forces, through specially designed openings at the outer edge of the disc. Depending on the number of revolutions of the disc and the wet product viscosity, particle sizes can range approximately from 5-250 pm [5.90-5.921. In spray atomizing of wet products, onecomponent nozzles give hollow spherical, free-flowing and relatively coarse products, while high or low pressure two-component nozzles, operated with propellants, give a fine product [5.93]. The mist droplets generated during atomizing are mixed with hot gas, which is centrally charged at the tower head. Some of the hot gas may be introduced tangentially to superimpose a rotation on the main flux of gas and product. In the cylindrical section of the spray tower, the product droplets sink downward in parallel or countercurrent flow to the gas. Droplets are therefore dried in seconds or fractions of a second. The obtained powder is discharged at the tower cone and fed to the conveying system by swinging lids, rotary-vane feeders, screw feeders, etc. A small amount of the fines (10-15%) reach the waste gas cleaning area with the waste gas, and is discharged there. With knowledge of the parameters influencing the particle size distribution of spray dried powders, and process experience with different products, new spray dryers with, for example, integrated fluidized beds are developed. Production of almost dust-free powders is now possible [5.110]. The height of the spray dry tower, Z is calculated from the sedimentation velocity w, of the droplet, the required drying time tg and the gas velocity wg.Referring to the largest droplet,

378

5 Drying

Fig. 5-37. System Buttner-Balfour spray drying unit. Representation according to Babcock-BSH AG, Krefeld. 1 Cooling gas fan 2 Gas distribution 3 Gas twisting baffles 4 Rotating spray disc 5 Filter the particle diameter, discharge velocity, 6 Gas fan spray angle, density ratio, drying time, etc., 7 Gas heater are given [5.94]. 8 Wet material vessel During the evaporation from the droplet 9 Spray tower 10 Exhaust gas discharge tube surface, which is generally the case for a 11 Cyclone droplet travelling in the spray dryer, the 12 Exhaust gas fan mass and heat transfer for in the gas distrib13 Dry material pneumatic conveyor uted droplets are [5.7] 14 Cyclone 15 Rotating valve

Nu = 2.0 + 0.6 . Re0.’

Sh 2 = tg * (w,k wg)

(5-127)

(+ applies for parallel flow, - for countercurrent flow of wet product and hot gas). For a quick estimate of the tower diameter and height, nomograms may be used if

= 2.0

+ 0.6

Re’.’

(5-128)

a

Sc0.33

(5 -129)

The characteristic length in the dimensionless number is the droplet diameter. The velocity is the relative velocity between the droplet and the drying gas (Eq. (5-125)). Table 5-10 presents some characteristic data for convection dryers.

0. Some characteristic data for convection drying.

5-15

free-flowing product, pulverized, pasty, lumpy

r

5-20

blocky, bulky, fibrous in large quantities, pasty product

r

5-18

large quantities of pellet-like to pasty products

yer

5000- 6000

6-20

agricultural products

6000-15000

0.15-1.5 for product gas overflow 0.1 - 12 for product gas through-flow

Specific heat requirement, water evaporation (kJ/kg)

Specific water evaporation rate (kg/(m2 . h))

e

Product condition

dryer pellet-like, pasty small quantities

3500-5000

4000-5000

4400-5000 in convection drying 3000-3800 in contact drying

Dryer type

- design data

Layer thickness on a kiln: 20-100 mm

Layer thickness: 150-200 mm Air velocity: 0.2-1 m/s Layer thickness: 20-50 mm Tunnel length: 20-60 m Tunnel width: 3-6 m Air velocity: 2-3 m/s 1-5 belts on top of each other Effective belt length: 5-50 m Effective belt width: 0.5-3 m Effective belt area: 3-100 m2 Belt velocity: 0.3-0.6 m/min Air velocity: 2-3 m/s Disc diameter: 2-5 m Disc area: 20-900 m2 Revolutions: 0.5-3 min-' Layer thickness: 10-20 mm Air velocity: = 2 m/s Power consumption: 6-25 kW

ions: c continuous, b batch

(continu

0. (continued) e

yer

r

ic

er

1.5-50 kg/(m3. h)

liquid (solution, suspension) to liquid-paste like products

25-500 kg/(m3 . h)

pulverized, crystalline, fine products, products able to be pneumatically conveyed

10-600

grainy, pulverized short-fibrous products, products able to be fluidized

4000- 6000

25-50 kg/(m3 . h) with direct heating up to 200 kg/(m3 . h)

nonsticky, lumpy pulverized or crystalline products, large throughput

Specific heat requirement, water evaporation (k J/kg)

Specific water evaporation rate (kg/(m2 . h))

Product condition

3500-6000

4000-5000

4000-6000

Dryer type - design data

Drum diameter: 1-4 m Ratio drum length/drum diameter: 4-7 Drum revolutions: 0.5-10 min-' Drum inclination: 1-6" Air velocity: 2-3 m/s Fluidized-bed area: 0.5-40 m2 Bed height: ca. 0.2-1.5 m

Riser diameter: 0.3-1.2 m Riser length: 15-30 m Superficial gas velocity: 10-40 m/s Particle size: 0.01 - 10 mm Drying time: 2-8 s Power consumption: 15-100 kW Tower diameter: up to =8.5 m Height: up to = 20 m Specific air necessary: 8-45 mh/kg water Power consumption: 0.1 -0.3 kW/kg water

5.11 Design of Dryers

5.11.2.9 Drum Dryer

In a drum dryer (rolling dryer or “adhesion-layer dryer”) [5.1, 5.2, 5.95, 5.961 a thin film of fluid or semifluid material (low viscosity to pulpy products) is evenly distributed by a spreader knife and retained on a drum, which slowly revolves with internal heating. The product is dried during one revolution of the drum (Fig. 5-38). Shortly before the drum dips again, the product is continuously scraped from the drum surface in the form of flakes or scales by means of scraper knives. Depending on the pressure and temperature in the drying room, the drying process is usually only contact or evaporation drying, or with the presence of a drying gas, a combination of contact and convection drying. According to the wettability, viscosity, and surface tension of the wet product, different charging devices should be used (Fig. 5-38). Drum dryers are manufactured with single or double drums as a single stage, or in series of two double drums with two stages. The residence time of the wet product in the drum t , is controlled by the revolutions n in such a way that it corresponds to the drying time, tg (5-130) where 7 is the angle between charging and discharging of the product on the drum.

5.11.2.10 Thin Film Evaporation Dryer (Vertical and Horizontal Dryer) The thin film evaporation dryer [5.28, 5.97, 5.981 enables concentration of soluted or suspended solids, to form a crystalline or powdery dry substance by evaporation of the solvent from a thin layer, under vacuum. The wet liquid product is fed over the

381

heating zone and, by means of a rotor, applied in a thin film to the heated wall. In the preheating and crystallization zone, the first crystals form and are thickened in the following slurry zone to a crystalline pulp. In the final powder zone the product is dried. Mounted swinging wipers prevent powder depositions and incrustation on the heating surface. Figure 5-39 shows schematically a thin-film evaporation dryer and a two stage thin-film evaporation drying unit. 5.11.2.11 Contact-Mixing Dryer With contact-mixing dryers [5.1, 5.21 the product is mechanically mixed, redistributed and intensively dried. Milling and scraping devices may be installed for lumpy products. Drying is usually carried out at atmospheric pressure or under vacuum conditions. In a plate dryer [5.29, 5.30, 5.461 (Fig. 5-29), without hot gas peripherals, free-flowing product is moved across heated discs using wiper arms, turnover blades, side scrapers, rakes, etc. The product is constantly overturned and alternately transferred from disc to disc through central openings or openings in the dryer walls and dried in this way. With a paddle dryer (Fig. 5-40) the liquid, pasty, or free-flowing wet product is evenly distributed by rotating paddles over the heated areas (jacket, hollow shaft, and paddles) and discharged at atmospheric pressure or under vacuum. Other contact-mixing dryer designs are based on the paddle dryer with the possibility of continuous operation. For example, the AP-dryer reactor [5.99, 5.1001 includes two parallel stirring gears, a main shaft, and a cleaning shaft. The main shaft is equipped with discs and mixing ingots. On the faster rotating cleaning shaft (counterrotation/synchronous) disc elements and knead ingots are also mounted. Due to the

382

5 Drying

I

f

W@

M

b

Fig. 5-38. Cylinder dryer, design Biittner. Representation according to Babcock-BSH AG, Krefeld.

I Single drum dryer 1 Heated drum 2 Vapor hood 3 Wet material trough 4 Roller for wet material 5 Charging roller 6 Scraper 7 Dry material discharge screw 11 Drum dryer arrangements.

Representation according to Krauss-Maffei AG, Munich. a) Feeding rolls for pasty materials b) Drum dryer arrangement for viscous materials and slurries c) Dip feeding for fluid materials d) Roller-feeding for fluid materials e) Spray-feeding by rotating sieve basket for thermally sensitive materials

5.11 Design of Dryers

c11:

383

Cooling water

cuurn pump

WF

d

VA

Sealing liquid

WF

Condenser

diL Raw product

t

Feed Pump

DM

llate drawal

Fig. 5-39. Thin-film evaporator-dryer. Representation according to Luwa - SMS Verfahrenstechnik, Zurich/Butzbach. a) Vertical thin-film evaporator-dryer b) Combined vertical and horizontal evaporator-dryer WF Wet material feed DM Dry material discharge VA Evaporated moisture (vapor)

384

Steam

5 Drying

I1 l l

Condensate

Cooling

water

Fig. 5-40. Paddle dryer. Representation according to Buss AG, Basel. BO Body RO Rotor with paddles DR Drive FR Frame VF Vapor filter VC Vapor condenser DD Discharge device VP Vacuum pump

special arrangement of discs and mixing ingots an intensive mixing and kneading effect is achieved as well as a 90% self cleaning of the heat transfer areas. On the shaft helically inclined welded mixing ingots cause an axial product transport. In a Discotherrn (operated continously or batchwise), [5.100], [5.101] interrupted disc elements with knead ingots or mixing hooks are mounted on the rotor shaft. In

t

A

the shell area not covered by the rotating kneading or mixing ingots, counterhooks are fixed at the shell. Due to the interaction of rotating kneading/mixing ingots and the fixed counterhooks an intensive mixing and kneading motion and a 90% self cleaning effect are achieved. In a Druvatherm, the wet product is flung, whirled and intensively mixed by shovels similar to plowshares. Between the shovels, agglomerates are destroyed by knives mounted on a head. Therefore, products showing a tendency to form lumps or which are normally difficult to dry can be processed. In a Drais dryer, which is designed similarly to the Druvatherm, the complete mixer may be removed in the axial direction for cleaning purposes.

5.11 Design of Dryers

385

5.11.2.12 Contact Dryer with Continuous Product Movement due to Gravity

(Fig. 5-41). The drying process is a batchwise operation.

Tumbler Dryer [5.30, 5.1021

Double Cone Dryer [5.102]

In a tumbler dryer, free-flowing products are dried under vacuum. Through a shift in the axis of rotation and the vessel axis, the jacket-heated drying vessel has a tumbling motion. The product inside the dryer follows this motion due to gravitational forces

In a double cone dryer (“mixing dryer”) a drying drum with the form of a double cone revolves around an axis of rotation, which is perpendicular to the drum axis. Bulk material with a poorer flow behavior is dried in batches, usually under vacuum conditions.

Fig. 5-41. Vacuum tumbler drying unit for polymer chips. Representation according to Babcock-BSH AG, Krefeld.

Tumbler dryer Scrubber Rotary piston pump Condenser Rotary gate valve pump Condensate receiver Heating and cooling agent unit

386

5 Drying

5.11 Design of Dryers

387

Contact Rotary -be Dryer (Contact Drum Dryer) [5.30]

acteristic data for a selection of contact dryer designs.

The rotary dryer in Fig. 5-30, with a jacket heated by flue gas, for example, is also suitable for free-flowing and dusty products. In a contact rotary tube dryer, the product is guided through the rotary drum by arranged tube bundles. The heat required for the drying is transferred from a heating medium flowing in the shell. Other contact dryers are described in detail in t5.1, 5.21. Table 5-11 gives some char-

5.11.3 Process Control of Dryers Process control of dryers is discussed, for example in [5.2, 5.103, 5.1041. Applications of process control computers to drying technology are described in [5.105]. A spray drying unit is presented in Fig. 5-42 as an example of the representation of process control installations and equipment.

4 Fig. 5-42. Flow and instrumentation diagram of a spray drying unit, pilot scale. Represented by Krauss-Maffei AG, Munich. 16 Mixed vessel 1 Fresh air filter with louvrc 17 Slurry dosing pump 2 Fresh air fan 18 Dosing pump 3 Hot gas generatiodgas burner 19 Filter 4 Hot gas distribution assembly 20 Homogenizer 5 Drying tower with hot-gas spiral 21 Spinning-disk atomizer distributor, sieve-tray top, tangential hot 22 Control device for spinning disk atomizer gas channel, air brush 23 Wet product ring line including lance 6 Dry product storage nozzles and atomizer pressure nozzle 7 Cyclone 8 Vibrated convcyer chute P Pressure 9 Venturi scrubber AP Pressure drop 10 Circulation pump T Temperature 1 I Compressed air sack filter F Flow 12 Main blower M Motor 13 Silencer I Indication 14 Fan C Control 15 Paste mixing bin A Alarm

388

5 Drying

Table 5-11. Characteristic data for a selection of contact dryer designs [5.1, 5.2, 5.45, 0.61. Dryer type

Product condition

Specific water evaporation rate (kg/(m2 * h))

Specific heat requirement for water evaporation (kJ/kg)

Cylinder dryer

liquid to pasty

15-75

3000 -4000

Thin-film dryer

liquid to pasty -muddy

50- 100

3000

Paddle dryer

liquid Pasty free-flowing

5-20

3000-3500

Drais-dryer

liquid Pasty free-flowing

10-30

3000

Contact rotary tube dryer

grainy free flowing

4-8

3200- 3600

Tumble dryer

granular free flowing

2- 10

3000

Abbreviations: c continuous, b batch

5.11 Design of Dryers

Operating mode

Design data

C

Cylinder diameter 0.6- 1.5 m Cylinder width: 0.8-3.6 m Effective cylinder circumfence: 75-90070 Cylinder revolutions: 2-20 min-' Heating area (per cylinder): 1.5-17 m2 Layer thickness: 0.1 - 1 mm Power consumption: 3-25 kW Diameter: 0.2-1.1 m Height: 2-12 m Heating area: 0.25-18 m2 Throughput: 10000 kg/h Driving power: 133 kW Diameter: 0.5-2.2 m Nominal volume: 1- 16 m3 Heating area: 1-60 m2 Construction length: 2-6 m Driving power: 1.5-50 kW

b, c

Content: 0.05-30 m3 Heating area: 0.5-60 m2 Driving power: 5-700 kW Throughput: 3600 kg/h

C

Drum diameter: 2-5 m Drum length: 4-8 m Drum inclination: 8- 15" Drum revolutions: 3-8 min-' Heating tube diameter: 100-130 mm Heating area: 200-3000 m2 Particle size: <20 mrn Power consumption: 2.4-40 kW

b

Drum volume: 0.1-50 m3 Heating area: 1.5-90 m2 Drum revolutions: 1- 12 min-' Power consumption: 0.2-40 kW

389

390

5 Drying

References [5.1] KRISCHER, O., KAST, W.: Vol. 1 Die wissenschaftlichen Grundlagen der Trocknungstechnik. Kroll, K.: Vol. 2, Trockner und Trocknungsverfahren. Aus der Reihe : Trocknungstechnik. SpringerVerlag, Berlin, Heidelberg 1978. [5.2] KNEULE,F. : Das Trocknen. Sauerlander, Aarau 1975. [5.3] KEEY,R. B.: Drying, Principles and Practice. Pergamon Press, Oxford 1972. [5.4] LYKOW,A. W. : Experimentelle und theoretische Grundlagen der Trocknung. Verlag Technik, Berlin 1955. [5.5] NONHEBEL, G . , and Moss, A. A. H.: Drying of Solids in the Chemical Industry. Butterworth, London 1971. [5.6] FIMNENKO, G. K., and LEBEDEW, P. D.: Einfiihrung in die Trockentechnik.Fachbuchverlag, Leipzig 1960. E. U. : Hochschulkursus [5.7] SCHLUNDER, Trocknungstechnik. TU Kar 1sr uhe. [5.8] KROLL,K.: Erjiahrenstechnik 5 (1971) 7, 279-288. [5.9] VORHOLZ,R. : Chem. Anl. VerJ 12 (1974) 44-50. [5.10] BAEHR,H. D.: Mollier i, x-Diagramm fiir feuchte Luft. Springer-Verlag, Berlin 1961. P. :Kaeltetech. Klim. 25 (1973) [5.1I] BERLINER, 3, 59-70. [5.12] HIEKE,W.: Wasserdampf und Luff in der Technik part 1. Steinadler, HiekeVerlag, Mannheirn 1980. W. : Kaeltetechnik 17 (1965), [5.13] HAUSSLER, 52-62. [5.14] T ~ E L E NP.: , Aujbereitungstechnik 10 (1969), 635. [5.15] KRISCHER,O., and MAHLER, K. : CIT 31 (1959) 2, 88-93. W.: Chem. Zng. Tech. 43 [5.16] SCHICKETANZ, (1971) 245-251. [5.17] BUHLMANN,S., PILHOFER,T., and Horss, J. : Verfahrenstechnik 6 (1972) 6, 196-202. [5.18] POERSCH, W., and ELE EN, P.: Aufbereitungstechnik 10 (1971), 610-621. E. U.: Chem. Ing. Tech. 48 [5.19] SCHLUNDER, (1976) 3, 190-198.

[5.20] KAST, W.: VDI Ber. 397 (1981) 1-8. W. : Zuckerindustrie 107 (1982), [5.21] POERSCH, 3, 195-204. F. Maschinen[5.22] APELT,J., and BAHNER, markt 88 (1982) 43, 872-875 and 70, 1410-1413. [5.23] APELT,J., and BAHNER, F.: Chem. Ind. 34 (1982) 6, 417-418. [5.24] GNIELINSKI, V. :Fortschr. Verfahrenstech. 22 (1984) Sec. C, 293-303. [5.25] KIMENOV,G. A.: ZFL 35 (1984) 3, 218-221. [5.26] BOSNJAKOVIC, F. : Technische Thermodynamik, 2 Vols. Verlag Steinkopff, Dresden 1965. E. W.: Elektrowaerme Int. 40 [5.27] MA", (1982) B 3/4, 177-181. [5.28] BACHMANN,R. : Verfahrenstechnik 6 (1972) 8, 269-274. [5.29] KESSLER, H. G.: Chem. Zng. Tech. 41 (1969) 7, 463-472. B.: Chem. Ing. Tech. 56 (1984) [5.30] VOSTEEN, 11, 858-859. [5.31] KOLLMANN, F., SCHNEIDER,A,, and BOHMER,G. : Forschungsber. 1689 Land Nordrhein Westfalen. [5.32] SINNING, B.: ~rfahrenstechnik7 (1973) 12, 372-375. W.: VDI Nachr. 14 (1979) [5.33] FIDTHMANN, 27. [5.34] ECKHARDT, L. J.: Forschungsber. 643, Land Nordrhein- Westfalen. [5.35] POPERT, F.: Elektrowaerme Int. 30 (1972) B 6, B 311-B 320. H.-C.: Siemens Energie[5.36] GRASSMANN, tech. 2 (1980) 7, 280-284. [5.37] JONES,P. L., and CROSS,A. D. : "Heat and Mass Transfer in a Radio-Frequency Dryer", Proc. of the Third International Drying Research Lim., Wolverhampton/ England 1982. H. : Wochenbl.Papierfabr. [5.38] GRASSMANN, 107 (1979) 17, 661-662 and DOKT 255/79. [5.39] Elektr. Verwalt: 51 (1976) 6, 148-151 and VTB 7629/18. [5.40] STEINBACH, G.: CZ. Chem. Tech. 2 (1973) 8, 323-327. [5.41] FISHER,R. R.: Freeze Drying of Foods. National Academy of Sciences, National Research Council, Washington D. C. 1962.

References [5.42] NEUMANN,K.: Grundriy der Gefriertrocknung. Musterschmidt, Gottingen 1955. [5.43] WILLEMER, H. VDZ Bildungswerk BW 6756 (1985) 1-11. [5.44] Company report: ,,Industrielles Gefriertrocknen," Fa. Leybold-Heraeus GmbH & Co. KG, Koln. [5.45] ,,Marktubersicht uber Trockner," Chem. Zng. Tech. 56 (1984) 11, A 585-A 603. B. : Chem. Ing. Tech. 56 (1984) [5.46] VOSTEEN, 11, 874-875. , 1. [5.47] ~ T T M A NE.NVerfahrenstechnik(1979) SCHILP,R., and BAUMGARTNER, S.: Verfahrenstechnik (1979) 7-8. NILL,€3. : Erfahrenstechnik (1980) 9. W.: Chem. Ing. Tech. 54 (1982)4, [5.48] FRITZ, 383-385. [5.49] MILOJEVIC,D. Z., and STEFANOVIC, M. S.: Chem. Eng. Commun. 13 (1982) 4-6, 261/269. W.: Chem. Zng. Tech. 47 (1975) [5.50] POERSCH, 18, 765. [5.51] BAUNACK, E : Aufbereitungstechnik 24 (1983) 4, 219-223, [5.52] SATTLER,K. : Thermische Trennverfahren. Aufgaben und Losungen, Auslegungsbeispiele. Vogel-Verlag, Wurzburg 1979. [5.53] LINDEMANN, W. : Aufbereitungstechnik 24 (1983) 4, 195-204. [5.54] REH, L.: Dissertation, T H Karlsruhe 1961. [5.55] REH, L.: Chem. Zng. Tech. 46 (1974) 5, 180- 189. [5.56] REH, L.: Chem. Ing. Tech. 49 (1977) 10, 786-795. [5.57] ERGUN,S.: Chem. Eng. Prog. 48 (1952), 89-94. [5.58] WERTHER, J.: Chem. Ing. Tech. 54 (1982) 10, 876-883. [5.59] MARTIN,H.: Chem. Ing. Tech. 52 (1980) 3, 199-209. [5.60] HEYDE,M., and KLOCKE,H.-J.: Verfahrenstechnik 13 (1979) 11, 886-892. J.: Chem. Zng. Tech. 56 (1984) [5.61] WERTHER, 3, 187-196. J.: Chem. Zng. Tech. 49 (1977) [5.62] WERTHER, 10, 777-785.

391

[5.63] MARTIN,H. : Chem. Ing. Tech. 56 (1984) 3 , 225-227. [5.64] MARTIN,H . : Chem. Zng. Tech. 54 (1982) 2, 156-157. L5.651 MERSMANN, A., NORTH,M., RINGER,D., and WUNDER,R.: Chem. Zng. Tech. 52 (1980) 3, 189-198. [5.66] POERSCH,W. : Aufbereitungstechnik 24 (1983) 4, 205-218. [5.67] POERSCH, W.: Maschinenmarkt 81 (1975) 88. [5.68] HOFFMANN,H., and MOLERUS,0.: Chem. Zng. Tech. 47 (1975) 23, MS 306/75. [5.69] DEMMICH, J., and BOHNET,M.: Verfahrenstechnik 12 (1978) 7 , 430-435. J.: Chem. Ing. Tech. 56 (1984) [5.70] DEMMICH, 3, 240-241. [5.71] WERTHER,J. : Aufbereitungstechnik 12 (1974) 670-677. [5.72] KESTNER,D.: Tech. Mitt. 78 (1985) 10, 489 - 497. [5.73] STEIN,W. A. : Chem. Ing. Tech. 45 (1973) 16, 1032-1039. [5.74] MARTIN,H., and SALEH,A. H.: Verfahrenstechnik 16 (1982) 3, 162-167. [5.75] BOHNET,M. : Chem. Ing. Tech. 55 (1983) 7 , 524-539. [5.76] WIRTH,K.-E. : Chem. Zng. Tech. 55 (1983) 2, 110-122. [5.77] GEIGER,A. : Verfahrenstechnik 2 (1968) 6 , 264-268. [5.78] SCHAUB,F.: Chem. Ing. Tech. 34 (1962) 3, 213-218. O., and HESS, D.: Maschi[5.79] ENSINGER, nenmarkt 81 (1975) 71. 1331-1332. [5.80] HEINZE,C.: Chem. Zng. Tech. 56 (1984) 3 , 238-239. [5.81] KOLLER, H. Chem. Ztg. 95 (1971)Tl-T3. M.: Aufbereitungs[5.82] BLUME,G., MATTER, technik 23 (1982) 4, 219-222. [5.83] RISENBY, C . : Verfahrenstechnik 16 (1982) 2, 88-91. [5.84] MASTERS, K.: Spray Drying. A n Introduction to Principles, Operational Practice and Applications. Leonard Hill Books, London 1972. [5.85] SCHUBERT,M., and VIEHWEG, H.: Spriihturmtechnik. Deutscher Verlag fur Grundstoffindustrie, Leipzig 1969.

392

5 Drying

[5.86] STEIN,W. A. : Chem. Ing. Tech. 44 (1972) 22, 1241- 1246. K.: Spray Drying Handbook. [5.87] MASTERS,

Verlag George Godwin Ltd., London 1979.

[5.88] HEIN, J.-C.,

RAFFLENBEUL, R., and BECKMANN, M.: Chem. Ing. Tech. 54

(1982) 9, 787-792. [5.89] LEE,D. A.: Chem. Eng. Prog. 77 (1981) 3, 34-38. E.: Chem. Ing. Tech. 45 [5.90] MEHRHARDT, (1973) 6, 401-402.

[5.91]BRAUER,H., and KRUGER,R.: Verfahrenstechnik 3 (1969) 3, 107-116. K., and SCHUSTER, I.: Ver[5.92] KLAMRCWH, fahrenstechnik 3 (1969) 3, 116-119. [5.93] HEGE,H.: Aujbereitungstechnik3 (1969)

142- 147. [5.94] HORTIG,H.-P.: Chem. Ing. Tech. 42 (1970) 6, 390-396. [5.95] MAHLER,K., and STOCKBURGER, D.: Chem. Ing. Tech. 37 (1965) 4, 406-414. [5.96] WYSOCKI, G.: Maschinenniarkt 79 (1973), 1710-1713. [5.97] WIDMER,F.: Chem. Ztg. 95 (1971) 18, 772-780. i5.981 WIDMER,F.: Vak. Tech. 18 (1969) 7, 147- 152. [5.99] LIST,H., and SCHWENK, W. : CZ Chem. Tech. 2 (1973) 11, 419-423. [5.100] Documents from Fa. List, Pratteln,

Switzerland.

[5.101] Documents from Fa. Krauss-Maffei AG,

Munchen.

[5.102] Documents from Fa. Henkhaus, Heu-

senstamm/Offenbach.

[5.103] AMME,K.: VDI Ber. 397 (1981) 39-42. [5.104] ROCK, H.: Regelungstech. Prax. 22 (1980) 11, 403-407. [5.105] WALLRAF, H.: Chem. Tech. Heidelberg 9 (1980) 12, 621-624. [5.106] MANN,E.: VDI Ber. 590 (1986) 55-71. [5.107] STUCHLY,S. S., and STUCHLY,M. A.: Adv. Drying 2 (1983) 53-71. t5.1081 GUPTA,R., and MUJUMDAR, A. S.: Adv. Drying 2 (1983) 155-192. [5.109] ALSTETTER,F.: Chem. Ing. Tech. 58 (1986) 6, 518-519. [5.110] HERBENER,R.: Chem. Ing. Tech. 59 (1987) 2, 112-117. [5.111] KUNII, D., and LEVENSPIEL, 0.: Fluid-

ization Engineering, 2nd Ed., Butterworth-Heinemann, USA 1991. [5.112] GELDART, D. : Gas Fluidization Technology, John Wiley & Sons Ltd., New York 1986. [5.113] TSOTSAS, E.,

GNIELINSKI, V., and SCHLUNDER, E. U.: “Drying of Solid Materials” Ullmann’s Encyclopedia of Industrial Chemistry. Vol. B2. VCH Verlagsgesellschaft, Weinheim 1988. J. F., [5.114] COULSON,J. M., RICHARDSON, BACKHURST, J. R., and HARKER,J. H . : “Drying” Chemical Engineering. Vol. 2. Pergamon Press, Elmsford, Oxford 1991.

Thermal Separation Processes: Principles and Design Klaus Sattler, Hans Jacob Feindt copyright 0VCH

Verlagsgesellschaft m h H , 1995

Extraction

6.1 Basic Concepts and Processes

solid by a solvent to form a solution. Figure 6-1 shows a simplified block diagram for a solid-liquid extraction process with solvent recovery. Solid-liquid extraction is mainly applied to extract metals from ores, oils from oil seeds, oil fruits and press cake, luxury items, and spices and pharmaceutical substances from plants and fruits. Low-boiling organic liquids and water are used as solvents. In high pressure extraction (HPE, distraction) compressed gas, usually in the supercritical or liquid state, is used as the solvent. Figure 6-2 shows a simplified block diagram for high pressure extraction by pressure and temperature swing, or by absorption or adsorption. HPE is also applied on a technical scale, mainly in the fields of:

Extraction is the selective dissolution, washing, or leaching of certain substances from a solid or liquid mixture by the aid of a liquid solvent. The valuable substance (or with a cleaning process, the pollutant) is transferred by a solvent as the receiving phase from the feed phase. Thus, extraction is a sorption process similar to adsorption and absorption, in which a selective auxiliary substance acts as the receiver phase. In contrast to other separation processes, extraction does not lead directly to pure key components. The feed phase after extraction (raffinate phase) contains not only the carrying substance which is inert to the solvent, but also solvent, and valuable key components. The receiving or extract phase, Solvent consists essentially of the solvent and key Solid f e e d L- - - - - - - I components. Regeneration of the solvent Preheating and purification of the key components requires additional separation processes, for example, rectification. It is important to intensively mix the prepared feed phase with the receiving solvent in Phase separation I Ifilter. centrifuge) extraction processes, and to allow sufficient contact time for the transfer of key component(s) between the phases. After the reRo f f i n o te E x t r o c t phose phase quired degree of mass transfer has been achieved, the raffinate and extract phases cation or evaporation/ must be separated. Mixing and phase separation must be repeated several times in an exVoluoble product traction unit to obtain good extraction yield. In solid-liquid f?XtrmtiOn (lemhkg) a Fig. 6-1. Solid-liquid extraction process certain component is dissolved out from the (leaching).

+;----

394

6 Extraction

a)

Gas feed used as solvent

Solvent (super critical or liquid state)

Crude feed

Extract phase

I

4 Rafinate phase I

Separation by pressure release or heating (solvent regeneration by pressure or temperature swing)

solvent

Valuable product

b)

Gas feed used as solvent

I

Crude feed

I

High pressure extraction

4 Raffinate phase

-1

Extract phase I

I

I

Solvent regeneration by sorption separation of the valuable component

Regenerated

I solvent Valuable product

Fig. 6-2. Simplified block diagram of a high pressure extraction. a) Pressure or temperature swing solvent regeneration. b) Sorptive solvent regeneration.

6.2 Liquid-Liquid Extraction

395

Food and luxury goods industry (e.g., extraction of vegetable fat and oil from oil seeds, production of hops and spices extracts, extraction of caffeine from coffee beans and nicotine from tobacco, by using hygienically safe supercritical gases as the solvent, such as carbon dioxide) Pharmaceutical industry Chemical and petrochemical industry (e. g., extraction of hydrocarbons from coal with ethylene or toluene as the solvent, deasphalting of crude oil with propane/propylene mixtures as the solvent agent)

phase, the raffinate (Fig. 6-3). The extract phase is then separated into the key component and the recycled solvent in an additional separation process, usually a rectification unit.

In liquid-liquid extraction (solvent extraction) two liquid phases take part. From a liquid feed consisting of a carrier phase T and the key component S, the key component is removed by means of a solvent L. The solvent, acting as an auxiliary substance, should essentially be insoluble in the carrier phase T, but the key component S should have good solubility in it. Mechanical phase separation at the end of the extraction process gives a solvent-rich product, the extract, and a purified carrier

Liquid-liquid extraction (LLE) is gaining more importance with the requirement of low energy consumption and is in competition with rectification processes. It is favored as a preseparation process if:

0

0 0

6.2 Liquid-Liquid Extraction 6.2.1 Fields of Application and Process Examples

0

0

The components to be separated are thermally sensitive or react with each other under higher temperature to produce unwanted side products The mixture components have a high or low boiling point so that rectification has

Liquid feed (carrier phase T+key component S )

Solvent L

Raffinate phase (carrier phase T+key component residue+ solvent residue)

Extract phase (solvent L i k e y component S+ carrier residue)

Key component S

Fig. 6-3. Liquid-liquid extraction, process schematic (solvent extraction).

396

6 Extraction

a (accomplished in distillation)

Fig. 6-4. Comparison of plant costs for distillation, extractive distillation and solvent extraction. Representation according to SOUDERS[6.10] and HAMPE[6.11]. tz = (.Y1/x,)/(.Y2/x2)

x , , y , molar fraction of component 1 in coexisting phases at equilibrium state (raffinate and extract

phase at equilibrium state).

0

0

0

to be carried out under a costly vacuum or under low temperature operation The boiling points of the mixture components are similar or an azeotrope is formed Several substances with a wide range of boiling points in a certain substance class are separated simultaneously from a liquid mixture (e. g., extraction of aromatics, see Fig. 6-5) The key component or pollutant is of low concentration in the mixture, therefore separation by distillation is very costly (economic application of LLE is, in the range of key component (or pollutant) concentrations from ca. 50-100 ppm to 10%)

Higher expenditure for LLE compared with distillation is worthwhile if a solvent of high selectivity, characterized by the separation factor a, is readily available. Invest-

ment costs for distillation, extractive distillation, and LLE are compared in Fig. 6-4 [6.10, 6.111. The separation factor i q a1,2 =

Yl ‘X2 ~

Y2 * Xl

where y l , y 2 , xl, x2 molar fraction of the lower (1) and higher (2) boiling component in the vapor and liquid phases (distillation), and in the extract and raffinate phases (LLE), respectively

For example, for a mixture containing components 1 and 2 to be separated by distillation, with a separation factor of a = 1.5, separation by extractive distillation or LLE is economic with separation factors of 2 or 6, respectively.

6.2 Liquid-Liquid Extraction Extract or

1.Stripper

f t

2 Stripper Aromotic substances Nonaromatic

Wash tower

Raffinate phase

1 c

397

Extract

1 Reflux

-

A

-

I

Water+solvent

I I

I I

Solvent+arorn&ic substances Solvent

A 7

Fig. 6-5. LURGI-Arosolvan process. Representation according to Lurgi GmbH, FrankfurUMain.

Some examples of practical applications for liquid-liquid extraction are listed in Table 6-1. Figure 6-5 shows a simplified flow diagram of a unit to extract aromatic substances using the Lma-Arosolvan process. An aromatic-rich feed is fed to the middle of the countercurrent tower extractor, operating according to the principles of a mixer-settler (see Chapter 6.2.4.1). A solvent consisting of an n-methyl pyrrollidoneglycol mixture is charged at the top. This flows as the heavy phase from the top to the bottom countercurrently to the light phase which flows in the opposite direction. Mixing and settling zones are passed in turns. Aromatic substances and some low-boiling substances in the feed are enriched in the extract phase, while the nonaromatic substances remain in the raffinate phase. The raffinate phase is discharged at the top and passed to a wash tower. The remaining carrier solvent is removed by washing with water. The extract phase is treated in two steps.

In the first stripping column, low-boiling nonaromatic substances and some benzol are distilled as the top product, and then fed back to the extractor as the lower reflux. In this stripping column, the water from the wash tower, loaded with solvent, is also regenerated and guided via the reflux vessel back to the wash tower. The bottom, or sump, product from the first stripping column, now free of nonaromatic substances, is separated in the second stripping column. Here, the top aromatic product is pure whilst the bottom product is aromatic-free solvent, which is fed back to the extractor. Figure 6-6 shows a flow diagram which explains this process. Figure. 6-7 shows the application of an LLE process to recover acetic acid from wastewater. Acetic acid is extracted from the feed by means of a solvent in an extraction column (ethyl acetate, methyl isobutylketone, etc.) leaving a final concentration of 0.1 -0.5 Yo. In a downstream stripping column, the solvent is removed from the gener-

398

6 Extraction

Table 6-1. Examples for commercial liquid-liquid extraction applications. Feed or process

Valuable substance or pollutant

Hydrocarbon mixtures like benzene reformate, hydrogenation raffinate from pyrolysis gasoline and oil gasification benzene [6.11] Udex (Dow-Chemical) Sulfolan (Shell) Arosolvan (Lurgi) Institut FranGais du Petrol Formex (SNAM-Progetti) Morphylan [6.1 21 (Krupp-Koppers) Mofex (Leuna-Werke) Crude oil fractions

Aromatic substance

Mercaptan

Sodium hydroxide solution etc.

Caprolactam synthesis mixture Tar distillate

Caprolactam

Benzene

Phenol

Sodium hydroxide solution etc.

Wastewater containing phenol Metallic salt solution [6.13]

Phenol

Petroleum naphtha Methylisobutyl ketone etc.

Metallic salt

Carboxylic acid etc. (if necessary reactive)

Acid residue in aqueous solution or wastewater Acetic acid recover [6.14]

Acid

Nitric acid recovery from wastewater Aqueous hydrofloric acid Nuclear process residues Base mixture for a reactive LLE [6.15] Metal ion extraction (Uran-plutonium extraction

Solvent

Diethylene glycol Triethylene glycol

+

+ Water

Tetrahydrothiophendioxide

N-Methylpyrrolidone t Water or Glycol + Dimethyl sulfoxide + Water N-Formylmorpholine + Water N-Formylmorpholine

+

Water

[Mono]methylformamide

Acetic acid Nitric acid Hydrofloric acid Rare-earth metals

Metal Uranium, plutonium

Ethyl acetate Methylisobutyl ketone etc. Secundary amine in kerosine Amine Organic solvent (if necessary reactive) Solvent combinations with reactive components Complex forming solvents Tributyl phosphate)

6.2 Liquid-Liquid Extraction

ated raffinate phase. The raffinate then leaves the unit. The top product from the stripper, a heteroazeotropic mixture, is condensed and separated in a separator. The extract (leaving the top of the extractor, a solvent-acetic acid mixture), is separated in a rectification column into acetic acid (bottom product) and solvent with some water content (top product). The water is then removed in a separator and is recycled to the stripper. The regenerated solvent is fed back to the extraction column.

Solvent recti;icationL-i

399

Stripper

Extractor F e e d m

om

1Acetic

water Make-up feed

Fig. 6-7. Acetic acid recovery by liquid-liquid 'eed Aromatic substance

nonaromatic

6.2.2 Solvent Requirements, Selection of Solvent

? , Arosolvc

Aromatic substance

extraction, simplified flow sheet. Representation according to [6.100].

Nonaromatic substance

Fig. 6-6. Flow-rate diagram of a LURGIArosolvan process. Representation according to Lurgi GmbH, Frankfurt/M.

The main problem in liquid-liquid extraction is the selection of the solvent. Of the different applicable solvents, a solvent characterized by an appropriately high loading capacity and a sufficient selectivity must be found, which also enables a minimum amount of recycled solvent possible in the LLE process. The loading capacity of the key component, or pollutant, to be transferred to the solvent gives the distribution of the component between the two liquid phases. This distribution equilibrium is described by the Nernst distribution law (Chapter 1.4.2.1). The amount of circulated solvent is a function of the loading capacity. The selectivity characterizes how much better the key component is extracted than the other components. The higher the selectivity of a solvent, the lower the number of separation stages required in the extractor. The solvent has to fulfill additional requirements besides capacity and selectivity.

400

6 Extraction

Table 6-2 gives an overview of the property requirements for the selection of a solvent. To evaluate the suitability of a solvent, the solvents and their properties are collected in a matrix (Fig. 6-8) [6.13]. The properties are then weighted for each solvent according to their importance and influence by using weighting factors qij, giving a weighted value qij wij for each property. The solvent with the highest sum of all the weighted properties is the optimal solvent.

0

0

0

0

As multistage extraction with phase cross flow As multistage extraction with phase countercurrent flow As multistage extraction with phase countercurrent flow and extract reflux As countercurrent distribution using two solvents

In the following sections the above variations are briefly discussed.

6.2.3.1 Single Stage Extraction

6.2.3 Liquid-Liquid Extraction Variations Liquid-liquid extraction is practically applied in the following variations: 0

As single stage extraction, also differential stagewise with recycled solvent (simplest and in laboratory-scale most common variation)

In a single stage extraction process the solution to be extracted, i. e., the feed F consisting of carrier substance T and the key component S (transfer component), is intensively mixed in the mixing section of the extraction device with the total amount of solvent La.After the distribution equilibrium is reached, the extract and raffinate phases leave the separation and settler zone of the extractor. The extract phase mainly consists

Table 6-2. Required solvent properties in LLE, aspects for solvent selection.

0 0

0

0

0 0 0 0 0

Sufficient loading capacity of key component per weight solvent (favorable distribution equilibrium, with chemically active solvents favorable distribution and reaction equilibrium). High selectivity. As low a miscibility as possible of the extraction solvent with the carrier. Easy and high separation of solvent from the extract phase (large difference of boiling points of solvent and key component, no formation of an azeotropic mixture). Large density difference of heavy and light phase to simplify phase separation and to avoid emulsion forming. Suitable interfacial tension (low surface tension favors formation of small droplets and promotes emulsion tendency; high surface tension causes formation of large droplets and thus lower mass transfer). Low viscosity (too high a viscosity reduces fluid mechanics and mass transfer properties). Low vapor pressure at operating temperature to avoid solvent losses by evaporation. Sufficient thermal and chemical stability. No corrosivity, no safety to use precaution and environmental impact. Good availability, low costs.

6.2 Liquid-Liquid Extraction

401

Properties Unit Weighting factor qij Solvent 1: Solvent 2:

Solvent n:

Weighted value q i z .wil Rating wi2 Weighted value q i 2 .wi2

Weighted value q;,,. win

* Marangoni instability: Instable interfacial convection due to locally different structure of the interface. Fig. 6-8. Solvent evaluation. Representation according to HAMPE[6.13].

of key component and solvent together with some of the carrier phase dissolved in the solvent. The raffinate phase is the residual liquid from which the solute or key component is removed, with some remaining solvent and key component (Fig. 6-9). If the extraction system only comprises the three components, carrier T, key component S , and solvent L , three balance equations are generally sufficient for an ex-

I

Fig. 6-9. Single-stage extraction. MX Mixer SE Settler, separator

tractor massbalance (see Chapter 1.3.1). With the nomenclature used in Fig. 6-9, a total molar balance gives La + F = R,

f

E,

(6-2)

A molar balance for the key component S gives La * y a+ F . X ,

= R,

*

X,

+ E,

‘y,

(6-3)

The molar balance is given with the carrier as the reference component

When the carrier and solvent are completely insoluble, the loadings X,Y (see also Table 1-4) are used in place of the fractions

402

6 Extraction

x,y. A balance for the key component, or

with the extraction factor

pollutant S referred to the now constant amount of the carrier T and solvent LT, gives T - X ,+ L,* Ya= T . X,

+ LT. Y ,

(6-5)

Hence, with known loading differences

X, - X , and Y, - Y, the solvent requirement LT follows.

The nomenclature used in Eqs. (6-2)-

(6-5) is:

F,L,,R,,E,

molar amount (or molar flow rates with continuous operation of the extractor) of the feed mixture, of the fed solvent from the regeneration stage, of the raffinate and the extract phases molar fractions in the raffinate and extract phases (the carrier has the index T) key component loadings of the raffinate and extract phase

X ,Y

x,y T=

F

~

1

, L,=-

+xu

where v is the solvent ratio referred to the feed carrier. The operating line for a single stage LLE extraction process, with distribution equilibrium, is shown in equilateral-triangular coordinates in Fig. 6-10 (Gibbs triangle). A general case is presented; the feed F, solvent L a , raffinate phase R, and extract phase E, may all contain all three components; the carrier, the key component or pollutant, and the solvent. All state points F, L a , R,, and E, lie inside the triangle. (If the feed contains no solvent, the state point F lies directly on the triangle side TS; if pure solvent is used, it has a state point at the corner L of the triangle). The state points R, and E, of the resulting phases both lie on S

La 1 + Y,

amount of carrier and solvent, respectively When the distribution equilibrium of the phases is reached

< T L TP T

I

Yw=K**X,

(6-6)

where K* is the distribution coefficient (see Chapter 1.4.2), referred to the loading. If the solvent used is free of the key component (Y, = 0), the loading of the raffinate phase X , follows from the balance equations, Eqs. (6-5) and (6-6)

x, = X ,

*

T T+K*.L,

1

=xu.-l + E

(6-7)

L

Fig. 6-10. Single-stage extraction process in equilateral triangular coordinates. OP One phase region TP Xvo phase region BC Binodal curve TL Tie line ML Mixing line F Feed La Solvent R, Raffinate E, Extract M Mixture point

6.2 Liquid-Liquid Extraction

the binodal curve and at the end points of the same tie line (see Chapter 1.4.2.2). A4 represents the mixture point of feed F and solvent La on the mixing line the course of which is determined by the Lever rule for the phases/mixture (see Chapter 1.4.2.2)

nu,

FM - La ML, F ~

~

(6-9)

During the exchange of the key component between the feed and solvent, the state points F and La move toward R, and E, until the distribution equilibrium is reached. A LLE is only possible if the mixing point A4 is in the two phase region, that is between A and B. Therefore, with a fixed feed, a minimum amount of solvent La,min is required. According to Eq. (6-9) La,min sets the mixing point A on the binodal curve. A mass balance for the key component in the mixing process with F and La,min, at the state point A gives the minimum required solvent ratio ~2,:~"

403

6.2.3.2 Differential Stagewise Extraction With differential stagewise (discontinuous) extraction, the solvent Lgand mixture F a r e intensively mixed in a vessel (Fig. 6-11). The raffinate R and extract E are formed at the end of the extraction. The extract phase is continuously withdrawn, and then separated by distillation into the key component S and the recycled solvent L , for example. In comparison with simple single stage extraction, the continuous reflux of solvent from the extract phase results in improved extraction of the key component from the mixture.

-I

1

E

(6-10) I

where x,, y, are the molar fractions of the key component in the feed and solvent, respectively, and X, is the molar fraction of the key component in the mixture at state point A. Single stage extraction may be operated continuously and discontinuously in agitated vessels with a separation section. With continuous operation, the agitated vessel used as the mixing zone has to be followed by a separator or settler (mixer and settler). As already shown, a single extraction stage may only act at best as a single theoretical stage if phase equilibrium is reached between the leaving raffinate and extract phases.

I

Fig. 6-11. Differential discontinuous extraction. EX Extractor ES Extract separation unit

6.2.3.3 Multistage Cross-Current Extraction Multistage cross-current extraction may be operated continuously or stagewise. In continuous cross-current extraction (Fig. 6-12), the mixture to be separated is fed into the first stage where it is treated with the solvent i,.From this extraction a raffinate Rl and an extract El are obtained.

404

6 Extraction

Fig. 6-12. Cross-current extraction. 1,2, . . ., i, . . ., n extraction stages, (n = Nt if equilibrium is reached in each stage).

The raffinate phase is guided from stage to stage, with fresh, regenerated solvent added to each stage. The extract of each stage is collected and processed. With stagewise extraction, the same process is carried out in an extraction vessel with agitator. A series of local extraction stages now become a timed sequence of mixing (and, therefore, extraction processes) and settling sequences, with withdrawal of the respective extract phase and addition of fresh solvent to the remaining raffinate phase. If the distribution equilibrium between the raffinate and extract phase is reached, with continuous and stagewise extraction in each extraction stage, the compositions of the raffinate phases R,, R,, . . ., R, and extract phases E,, E,, ..., En are obtained from the state points in the Gibb's triangle (Fig. 6-13), by the method explained in Fig. 6-10. The number of state points for the raffinate or extract phases then corresponds to the number of the required theoretical stages N1 in the cross-flow cascade. In each stage, the key component is extracted from the feed with an initial concentration x,, state point F to a final concentration xw, according to point R , = R,.

S

T

L

Fig. 6-13. Raffinate and extract compositions in cross-current extraction. BC Binodal curve F Feed flow state point F . . . La Solvent flow state point L , , L,, L3 R,,R,, R, = R, Raffinate flow state point R , , R 2 , R 3= Rw E l ,E,, E, = E, Extract flow state point EI,E,,E3 E , M , , M,, M3 Mixture points r i , , ri2, ri, of raffinate and solvent -.-.-. Tie line _____ Mixing line

405

6.2 Liquid-Liquid Extraction

A total mass balance for the key component and the solvent over the ith stage of the cross-flow cascade (Fig. 6-12) gives

R i P l + Li = ri,

(6-11)

where ri.I = Rl. + E l.

(6-17)

The raffinate flow rate Ri and the extract flow rate Ei are then (6-18) .

where raffinate flow rate to ith stage Ri-1, molar flow rate of solvent (reL ; , L,; ceiving phase) or pure solvent to the ith stage xi - 1 , xL, - molar fraction of the key component and solvent in Ri Yct molar fraction of the key component in Li x ~ , , , x ~ ,molar ~ , , fraction of the key component and solvent in the mixture flux rii of the ith stage

E.=

rii

*

(XM,i

-Xi)

(6-19)

Yi - xi

(If for each extraction stage, the solvent ratio v J , = ~ L i / R i- 1 is fixed, the mixing point M , for the mixture ri, is given as the ratio Ri_,Mi/MiL,= v&, on the line R,-,L,. Hence, with equilibrium reached in stage i, the state points Ei and Ri for the extract phase Ei and the raffinate phase R, are the corresponding ends of the tie lines, intersecting Mi, with the binodal curve). The procedure for cross-flow extraction may be substantially simplified by neglect(The molar fractions of the fed carrier re- ing the low solubility of the carrier and solsult from the stoichiometric summation vent. Therefore, the carrier flow rate T is condition Ex, = 1 and E y, = 1 for each re- constant throughout the entire cascade and spective mixture). the solvent flux LTiis constant in each sinFrom Eqs. (6-11)-(6-13), the coordinates gle stage i. The concentration of the key of the respective mixing point Mi as xM,,, component may be given as a loading rexM,L,i and 1 - xM,,- x ~ ,follow ~ , ~ ferred to the carrier or solvent. A mass balance for the key component over the ithex(6-14) traction stage then gives

,

~~

T.xi-l+ L , ; -

and

Y,= T . X , + L , ; *

r, (6-20)

xM,L,i =

Ri-1 . x ~ , i - l +L,i Ri-1 + Li

(6-15)

The mixture hi disintegrates into the flow rates R j and E; with a key component molar fiaction of xi and yi, leaving the extraction stage i. Therefore,

+

R, . xi E; * y ; = ri,

XM,i

(6-16)

From this, the amount of pure solvent required for the ith stage follows

L T i=

T . (q, -4) Y.- Yu

(6-21)

- I

if all solvent flow rates L , , . . ., L , . . . L, have the same initial load Y, of the key component. X i and X, are the key component

406

6 Extraction

loadings for the carrier flow T entering and leaving stage i. The required number of theoretical stages Nt of the cross-flow cascade is obtained by graphical means from the operating diagram (Fig. 6-14). Nt corresponds to the number of intersections of each stage with the operating line and equilibrium curve starting from point A with the coordinates X , and Y,, and then ending at the required final load of the raffinate phase X, that is withdrawn from the last stage. Individual balance lines are characterized by slopes according to Eq. (6-21) tan xi =

--

5 - Yu

x,-,-x,

-

(If the distribution equilibrium is not reached in individual stages of the crossflow cascade, the described method may also be applied to find the actual or practical number of stages Np.An auxiliary line has to be introduced in Fig. 6-14 instead of the equilibrium line, which considers the enrichment ratio for each stage). The solvent requirement for the entire cross-flow cascade is

_ _ T_ _- - _1 LTi

vi

(6-22) which are fixed by the chosen solvent ratio vi (referred to the carrier).

where Xu and X, are the key component loadings of the feed F and the raffinate Rm withdrawn from the last extraction stage, Yu is the key component loading of the solvent, and Y,,m is the mean loading of all extract flow rates B1, . . ., Ei, E,. If the solvent is free of key component, Y, = 0, and the solvent flux to each stage is the same L,, = ... = L r i = ... = L T n , the loading ratio is then

-- -

xu

1

(1 + &)"

(6-24)

where Nt is the number of theoretical extraction stages and

I

Fig. 6-14. Operating diagram of a cross-current extraction cascade with three stages. EC Equilibrium curve BL 1, BL2,BL3 Balance line of stages 1, 2, 3 tanx, = -T/LTl Slope of each balance line Y Solvent load with key component X Carrier phase load with key component

(6-25)

is the extraction factor, equal for all stages. Eq. (6-24) may be also applied to differential stagewise extraction operated as a stagewise cross-flow extraction process, with the number of extraction cycles tending to infinity (n -,03). The extract phase may be withdrawn continuously and fresh solvent is fed. For N,4 03,

6.2 Liquid-Liquid Extraction

(6-26)

where Eg =

K* . L,,

T

= ln-

XCI

Xul

(6-27)

is the overall extraction factor, and L,, the total amount of solvent required. E,

6.2.3.4 Multistage Countercurrent Extraction With countercurrent extraction the mixture being treated flows countercurrently against the solvent through several extraction vessels in series (Mixer-Settlers) or more commonly, through extraction columns (Fig. 6-15). The feed F and solvent La flow from opposite ends in a multistage extractor toward each other, hence the raffinate flow R,,, comes into contact with fresh or regenerated solvent La and the extract flow Em comes into contact with the feed l? This leads to high key component concentration gradients between the entering and leaving phases. Therefore, good enrichment in the extract phase and considerable purification of the key component in the carrier phase are possible. In practice, columns operated with countercurrent extraction often have- the heavy phase charged at the top and the light phase at the bottom. One of the two phases is divided into droplets either once or stagewise, and then moves in the form of a droplet swarm (disperse phase) through the continuous (coherent) second phase. To ensure high phase turbulence and, therefore, large interfacial areas for good key component (or transfer component) mass transfer between the phases, columns contain internals such as filling materials, packing, sieve trays, and agitators. Energy dissipation is caused by pulsation and/or using rotating stirring elements.

407

Choice of the Disperse Phase

Basically both phases, the feed and solvent, may be used as the dispersed phase. However, it must be noted that in general, only one of the two phases is favorably dispersed with respect to mass transfer and phase flow. The following criteria may rule the selection, although they may also lead to the reverse decision in some cases, and therefore a compromise is necessary: The key component should be transferred from the continuous phase to the disperse phase To achieve the maximum mass transfer area, the phase with the larger volumetric flow rate should be dispersed With view to maximum throughput, the more viscous phase should be dispersed The disperse phase should have the smaller surface tension so it can be more easily broken down Assuming a holdup of ca. 10-30% for the disperse phase, which is hence considerably smaller than that of the continuous phase, the more expensive, explosive, or toxic substance should be dispersed To achieve small droplets and hence a larger transfer area, the component with the poorer wettability should be dispersed by the dispersion device If the physical and volumetric properties of the phase drastically change d u e t o the mass transfer, it may be convenient to move the interface between the two phases to the middle of the column. The light phase is then dispersed in the bottom section and the heavy phase is dispersed in the top section of the column. Mass Balance, Stage Concept

A mass balance for the key component over the total column with the nomenclature

408

6 Extraction

b’

6A

E t X,I

I I

I

I

I

I I I

RXX-

/TI

r-

‘‘‘k,LToY

I

I

I I I

I

I

HP

Fig. 6-15. Countercurrent extraction. a) Countercurrent extraction in a mixer-settler cascade. b) Countercurrent extraction in a column. c) Dispersed and continuous phase in a countercurrent extraction column.

LP HP LIC IP BA

Light phase, in this case dispersed Heavy phase, here the continous phase Level control Interphase Balance area

6.2 Liquid-Liquid Extraction

This is the equation of the balance or operating line in the operating diagram in Fig. 6-17.

Total cost

I

409

The solvent ratio v L7T

V=-

I

I

is of similar importance for extraction as is the reflux ratio for rectification, and the solvent ratio for absorption. The slope of the balance line is fixed by the solvent ratio v, and may not exceed the minimum solvent ratio vmin for a certain extraction problem. vminis graphically determined by the slope of the balance line which intersects the points A and S, or calculated by

cost Y-

Yu.opt

Total cost

v min . =

Investment cost I

nin

Vopt

v-

Fig. 6-16. Determination of the optimal remaining solvent concentration after regeneration (a) and the optimal solvent ratio (b) in extraction columns. Y, Remaining load after solvent regeneration v Solvent ratio

From a mass balance for the key component in the top section of the extraction column (Fig. 6-15) the relationship between the loading X of the carrier and the loading Y of the solvent in any column cross-sectional area is

T Y=,*X+

=,

T Y,--.X,

=,

(6-30)

(6-29)

y,,rnax -

(6-31)

r,

.

Yl,,,,

f1

x a x u J

yw

1

ya

Fig. 6-17. Operating diagram for constant flow rates of carrier and solvent in countercurrent extraction. EC Equilibrium curve B L l Balance line for vmin BL2 Balance line for v > vmin Y Solvent load with key component X Carrier load with key component

410

6 Extraction

vmin implies that the feed and the leaving extract phase are in phase equilibrium. This is only possible with an infinite number of separating stages or with an infinite transfer time. Therefore, the actual solvent ratio v has to be chosen such that it is larger than vmin. The value of v for extractor design and operation should result from cost estimates according to Fig. 6-16. With the known course of the equilibrium curve for fixed extraction conditions, the required number of theoretical stages Nt is determined by drawing steps between the equilibrium curve and the balance line, analogous to the McCabe-Thiele method (see Chapter 2.5.2.4 and Fig. 6-18). If the distribution coefficient K* is constant in the loading range of interest X u 2 X 2 X,, the equilibrium curve is linear in the operating diagram. The required number of theoretical stages Nt is then calculated from

/Ec

Fig. 6-18. Graphical determination of required number of theoretical stages. EC Equilibrium curve BL Balance line Y Solvent load with key component X Carrier load with key component

If two fluxes, characterized by the state points A and B, are mixed, the state point of the mixture M, lies on the connecting line between A and B. However, if from Nt = -1 flow rate mA a part mBis removed, the state In E point M2 of the remaining mixture mA - rhB (6-32) also lies on the connecting line AB,but instead lies outside A and B on an extension where E is the extraction factor (see Chapter 1.4.2). Application of Eqs. (6-34)-(6-36) im(6-33) plies: the state points of the fictitious mixtures b must lie on an extension of the connecting line between the state points F and In the case when the carrier T and solvent E, or R , and La or R and E, respectively. L are considerably miscible, countercurrent All three fictitious differential flow rates D extraction calculations have to be carried have a common state point which is named out using triangular coordinates (Gibbs tri- the “pol” P, and is the intersection of the angle). Molar balances over the top section extended lines FE, and R,,,if the state of the column and over the total column, points F of the feed mixture 8 E, of the extract phase E,, R, of the raffinate phase with the nomenclature in Fig. 6-15, give A , and L, of the solvent phase L , are . . F - E, = fi - E = D (6-34) given. P gives the fictitious composition of the two phase system for fixed flow rates of (6-35) R and 8,and remains in the same place 6, - L o = R - E = D during the extraction process. The pol loca. . (6-36) tion is fixed by the feedkolvent ratio F - E , = R, - L , = D

6.2 Liquid-Liquid Extraction

(6-37)

According to a graphical method by HUNTERand NASH [6.17], the required number of theoretical separation stages Nt for the distribution of the key component between the extract phase E , and the raffinate phase R,, may then be determined using the Gibbs triangle. One has only to consider that the state points of phases in equilibrium leaving the respective stages lie on the binodal curve and at the end of the common tie lines of phases in contact in a common cross section lie on the binodal curve and on a common pol or cross-section line The Hunter-Nash method is illustrated in Fig. 6-19. The state points F and E, are connected by the top pol line. F is the state point of the mixture entering the top sepa-

A

411

ration stage. E, is the state point of the extract phase leaving the top stage E,. If the top stage acts as one theoretical separation stage, the leaving phases must be in phase equilibrium. R, is the state point of the raffinate phase 6,and lies on the binodal curve at the intersection with the tie line through E,. The raffinate phase 6,and extract phase come into contact in the cross-section below the top stage. State point El of El has to lie on the binodal curve and the pol line passing through R,, and is therefore known. When selecting the state points R,, R,, . . . and E,, E,, . . ., according to the method as described for the top stage, the number of stages is found from the number of the raffinate state points in the diagram. The minimum solvent ratio is referred to the feed, p =

Lmin ~

F

=,*.

L T min =A

F

(6-38)

and is derived from Fig. 6-20 by the Lever rule for the phases, with the assumption of phase equilibrium between the leaving extract phase E, and the feed I? Therefore, the state point of the extract phase Em,, will then lie on a tie line passing through F.

T

P

Fig. 6-19. HUNTER and NASHmethod to graphically determine the number of theoretical stages of

countercurrent extraction units. BC Binodal curve TL

PL

Tie line Pol line

L Pure solvent (triangular corner) F, E,, R,, La State points of feed and leaving phases

412

6 Extraction

S

tray efficiency Ed. If there is plug flow then the relationship between &d and Ed is (6-39.2) where

L

P

Fig. 6-20. Determination of minimum solvent ratio vkin. BC Binodal curve TL Tie line PL Pol line

(6-39)

E

is the extraction factor.

Generally, the mass transfer between two liquid phases in liquid-liquid extraction is substantially slower than that for rectification processes. Small values of the tray efficiency factors therefore follow. With average extraction, tray efficiency factors range from 0.3-0.7 with sieve trays, and 0.3-0.6 with Koch cascade trays. Only in mixer-settler cascades and with centrifugal extractors is the actual tray efficiency almost the theoretical tray efficiency. Introducing the concept of height of a theoretical separation stage, HETS to evaluate the actual height of an extraction column giving the same effect, the total height of the column Z for mass transfer is

(Since the initial load Y, is usually very small, imin is approximately equal to &,,in. The minimum solvent ratio v A * , ~ ~ ~ corresponds to the minimum solvent ratio Z = N t .HETS (6-40) referred to the feed, ~4,). To determine the actual number of re- For the determination of the separation efquired stages Np for the extraction column, ficiency which is characterized by the enthe number of theoretical stages Nt and the richment ratio or HETS in the case of liqstage efficiency factor EMd (MURPHREE uid-liquid extraction, results from experistage efficiency factor) have to be known. ments on a pilot-plant scale, are required to For the case of mass transfer from the con- a greater extent than in rectification. tinuous to the disperse phase with respect to The separation effect in extraction prothe dispersed phase, EMdis cesses is not only a function of common variables such as the phase column loads, Y n -Yn -1 but also of variables which depend on the (6-39.1) EMd = internals, the concentration profile along Y*(xn)- Y n - 1 the column, the substance compositions, where yn, y n - , are the concentrations of and substance properties. These additional the key component in the disperse phase variables and their influence make scale-up above and below the stage of interest, re- to production units, based on the results spectively, and y*(xn)is the stage exit con- found in pilot-scale plants, questionable. centration in equilibrium with the concen- The variables specific to extraction are tration of the continuous phase x,. For the listed in Table 6-3. Besides the mass transfer coefficient, the limiting case of ideal mixing in the separation stage, EMd corresponds with the local separation effect is mainly influenced by

6.2 Liquid-Liquid Extraction

413

Table 6-3. Specific extraction variables influencing the separation effect (exchange ratio, height of a theoretical separation stage HETS, height of a transfer unit HTU). ~

~~

Variable

Effect

Fluid mechanics Droplet formation Droplet size, droplet size distribution Droplet swarm motion Droplet coalescence

Small droplets and a narrow droplet size distribution favor mass transfer

Interfacial effects Interfacial tension gradient Viscosity gradient Marangoni effect (near interfacial liquid layers move from area of low interfacial tension to areas of higher interfacial tension)

Interfacial turbulences caused by interfacial tension gradient increase the mass transfer to the interphase compared to the mass transfer by only diffusion

Mixing effects Longitudinal mixing, back mixing (axial carriage of smallest droplets by the continuous phase, continous phase dragged by droplets in the wake) Cross mixing (radial mixing caused by cross flow diffusion)

Compared to the ideal case of plug flow the driving force for mass transfer is reduced by mixing effects

the mass transfer surface, which is directly proportional to the holdup of the dispersed phase and inversely proportional to the droplet diameter. Therefore, the effects of coalescence have a large impact on the separation effect. Stage-to-stage calculations are easy to do if the miscibility of the carrier and solvent may be considered as being negligible. For example, a molar balance for the key component over a three-stage countercurrent extractor (Fig. 6-21) gives the equation system: 1. Stage

-

T * X , + LT Y2= T X I

-I-L ,

*

Y,

(6-41)

References

3. Stage T . X2

+ LT

*

Y, = T * X3 + L

-3

(6-43)

In the case of distribution equilibrium in each stage, (6-44)

Introducing the solvent ratio, v = L T / T and the extraction factor E = K* . v, a system of three linear equations is finally achieved, with the unknowns X I , X 2 , and X 3 , and, therefore, q , &, and &

- (1 + C) . X I + E * Xz =

-X ,

X ~ - ( l + & ) . X 2 + & * X=3o

X , - (1

+ E)

*

(6-45) (6-46)

X,= - v Y, (6-47) *

414

6 Extraction

D=

1

-(1

+ &) 1 0

-(1

E

+ &) 1

-(1

+ &) (6-49)

XI

\'

L

is the coefficient determinant, DXi is the determinant for the respective unknown X i , following from D.In the equation system Eqs. (6-45)-(6-47) the coefficient of Xi is replaced by the components of the solution vector

yz

( -5) - v * Y,

For any number of stages the coefficient determinant D and the solution vector can be appropriately extended. If the initial solvent load is negligible, Y, = 0, the calculation is then considerably simplified. For any number of stages, n 3 N,,

3. Stage

xu-

x,

~-

Fig. 6-21. Material balance of a three-stage countercurrent extractor.

E-1

En+'-

1

(6-50)

From Fig. 6-22, based on Eq. (6-50), the required number of theoretical stages N,may be derived for an extraction yield 7 (Yo), if the extraction factor E is given. (6-51)

X I ,X , and X , may be calculated by following the Cramer rule for this equation system,

(6-48) if the entry loads X , , Y,, the solvent ratio v, the distribution coefficient K*, and the extraction factor E are given, where

HTU, NTU Concept Using the concept discussed in Chapters 1.7 and 1.9 for LLE, the height required for mass transfer Z for a countercurrent extraction column is

-

6.2 Liquid-Liquid Extraction

415

100 -

I

Jt

10

-

1

I

-

I

99.9

99

7[%1

I

I

90

0

Fig. 6-22. Determination of extraction yield. Representation according to BRUNNERand Westfalia Separator AG, Oelde [6.20]. Nt Number of theoretical separation stages q Extraction yield

= HTUo,.

NTUoE

(6-52)

where cross-sectional area of the extraction column kE,kR overall mass transfer coefficient referred to the extract and raffinate phases

AQ

or

where the key component loadings X and Y of the carrier and solvent, respectively, are referred to as the concentration scale

-= - + kR PR

BE,PR

(6-55) DE

mass transfer coefficient, referred to extract and raffinate phases

416

6 Extraction

effective specific volumetric mass transfer area (phase boundary area) loadings of the raffinate and exX,Y tract phases by the key component in one reference cross-section of the column (see Fig. 6-18) X*, Y* phase equilibrium loadings referred to X and Y (see Fig. 6-18) HTU,,, height of the transfer unit referred to extract and HTU,, raffinate phases, respectively NTU,,, number of the transfer units referred to extract and NTU,, raffinate phases, respectively ae

The heights of a transfer unit HTU,, and HTU,, are found empirically under conditions as close as possible to the operating conditions from the product of k E .a, and k, . a,. Transformation of the results obtained in an experimental column to a technical column is only possible based on scafe-i.rp relations~ips found empirically. For example, according to BAUER[6.21], for agitated columns

(2)

1/3

HTU, = HTU, ’

(5-56)

and according to TRORNTHON [6.22] for pulsed tray columns, HTU, = HTU,,. exp[C. (d, - d,)]

(6-57)

where d is the column diameter and the index t refers to the technical column and v to the experimental column. The constant C in Eq. (6-57) is, according to REISSINGERet al. [6.23], 1.64 for column diameters up to 300 mm. For larger diameters, C has to be reduced. Generally, no measurement of the concentration profile for the determination of the operating line Y ( X )is transferable. The balance line in Fig. 6-18 must be used to determine NTU, but its form assumes

plug flow in the column. The evaluation of the integral expressions in Eqs. (6-52) and (6-53), follows the method presented in Chapter 1.9.4. When small columns and technical columns are compared, it is observed that, for a given extraction effect, a column with a larger diameter requires a larger height for mass transfer Z,,,. This is due to axial and radial mixing effects which cause deviations from a uniform, fluid particle residence time distribution in both the disperse phase and the continuous phase. The driving forces are reduced by this mixing, compared with the ideal case of plug flow. The mixing effects are caused by: 0

0

0

0

Molecular and turbulent diffusion in axial and radial directions Carrying of continuous phase in the wake of a droplet Carrying of the smallest droplets by the continuous phase Cross flow

The negative influence increases with increasing column diameter. Figure 6-23 shows the concentration profile of both phases along the column height z for the cases of ideal plug flow and pronounced longitudinal mixing. With longitudinal mixing a distinct increase in concentration occurs at the point where the phases enter, due to mixing of the column contents. Mixing therefore causes a partial equalization of the concentrations along the extraction column length and hence a decrease in the driving force. To compensate this effect, a larger height for mass transfer, Z,,, is required compared with the height Z found by a plug flow model. Eqs. (6-52) and (6-53) are used to calculate Zefl, in which either HTU or NTU have to corrected ( T U , NTU) to consider the mixing, so

417

6.2 Liquid-Liquid Extraction

a)

Z

1

EC

-0 0 0

a

t

VI

0

L

a

Z

c

V

2

c

X

W

0

y,

x, Y

-

XU

Xa Raffinate phase load

X-

Fig. 6-23. Influence of axial mixing on concentration curves (a) and number of separation stages (b). a) --- Concentration curve at plug flow

- Concentration curve at pronounced mixing

b) EC

OL BL Y X

Equilibrium curve Operating line at pronounced mixing Balance line at plug flow Extract phase load Raffinate phase load

or

Z,,,

0

= HTU.NTU

(6-59)

From Eq. (6-58), the apparent height of a transfer unit is

HTU = HTU f HDU

(6-60)

when the mixing effect is considered. According Eq. (6-60), F U is the sum of the height of a transfer unit HTU for plug flow and the height of a dispersion unit HDU. Mixing effects may be found experimentally, by measuring the residence time distribution of the phases. Mixing flow models are used to convert the results into half empirical correlations (see, for example, [0.4]):

0

The cell model (stage model) describes hydrodynamic flow in countercurrent columns analogously to the flow in a number of equal, ideally mixed vessels, as individual mixing stages (cells). Between the individual mixing stages, no mixing takes place. The number of cells is the adjustable parameter for the measured residence time distribution in the mixing cell cascade. Stronger backmixing leads to a smaller number of cells, or to a larger cell height. In the more complicated backfiow model, where the column is also divided into individual mixing cells, as in the cell model, additional longitudinal mixing between the cells in form of “mixing flow” is included.

418

6 Extraction

The dispersion model (diffusion model) explains the deviation of the real flow profile from the ideal plug flow profile due to dispersion analogously to molecular diffusion. For example, in the continuous phase, the axial distribution of the key component is (6-61)

where dispersion molar flow rate of the key component in the continuous phase (current density) density and molar mass of the continuous phase concentration gradient of the key component axial diffusion coefficient in the continuous phase

nD, c

ec9Mc

dx/dz

D ,C

The Bodenstein number Bo is the ratio of pure convective axial mass transfer and mass transfer due to longitudinal mixing, and is a measure of the substance dispersion in the real column flow and the mixing effects. Large values of Bo characterize narrow, small values but strong scattered residence time spectra. Bod and Bo,. must be known to calculate HDU. If the extraction factor E = 1, which is true for many technical applications, the apparent height of a transfer unit HTU considering Eqs. (6-62) and (6-63) is __

HTU = HTU + ___ wc

Du~,c

HDU

-

1

L,.

-+-

Bod

0.8 lne + L , . E I.,.(&- 1)

(6-64)

0.33

(6-65)

(6-62) (%)?

BO,.

where Bod and Bo, are the Bodenstein numbers ( E extraction factor, Lc characteristic length, height, wd, w,. flow velocities of the dispersed and continuous phases)

wd

For example, to calculate the dispersion coefficient for a Rotating Disc Contactor (RDC), given in [6.24],

Calculation of the axial flow dispersion assumes knowledge of the dispersion coefficients for the continuous and dispersed phases, as a function of their variables. The dispersion model is valid if the mixing effects are small. Characteristic parameters of the dispersion and cell model are interchangeable. The height of a dispersion unit HDU is calculated according to ZEMERDINGand ZUIDERWEG I6.241, [6.25] 1

Dax,d + -_

where

I*):(

[(%)? -

(6-66)

P

dispersed phase fraction (holdup) Re,,Re, Reynolds numbers of the stirrer and the continuous phase

Re,, =

d; . n ~

VC

, Re,.=

d

*

W,

__ VC

(6-67)

w, = n . d, . n circumferential velocity of wd, wc,

d, d,,d,,

zz

n

the stirrer flow velocity of the dispersed and continuous phases, number of stirrer revolutions diameters of the column and impeller and the stator ring height of mixing cell

Application of the one-dimensional dispersion model gives the differential equations ~0.41

(6-68)

d=

1/-

v$

6.2 Liquid-Liquid Extraction

4.r;;

71

wi

*

= 1.128

419

(6-70)

The maximum phase velocity wi, similar to the other countercurrent flow columns, is related to the phase velocity at the flooding point wif wi

(0.5

... 0.8)

*

Wif

(6-71)

Most of the concepts used to calculate two phase flow in extraction columns are based on a simple model of two countercurrent liquid phases (the two phase model of GAYLER,ROBERTS,and PRATT[6.27]). The relative velocity w,, between the phases is the sum of the effective phase velocities (6-72)

(6-69) With known overall mass transfer coefficients kd and kc, specific volumetric interface area a, and axial dispersion coefficient, Duxthe solution gives the actual concentration profile of the key component in the column. In [6.26], methods to measure the longitudinal mixing in countercurrent extraction columns are described and approaches to calculate the Bodenstein number and the axial dispersion coefficient for common extractor designs are given.

Column Diameter, Holdup, Interface Area The diameter of an extraction column d is determined by the throughput of one phase i (extract or raffinate phase) and the maximum superficial velocity or volumetric flux, w iper unit cross-sectional area of the column

where w d is the velocity of the disperse phase, w, the velocity of the continuous phase, and y) is the holdup of the dispersed phase. A reduction factor 6 has to be introduced for columns with internals reducing the cross-sectional area (Eq. (6-72)). This is the void fraction of packed columns; introduc ing the relative void fraction it is (6-73 According to ~ O R N T O N[6.28] the relative velocity w,, is a linear function of (1 - 9). Extrapolation of the holdup for a single droplet ( y ) 0) gives the characteristic velocity of a droplet J +

WreI

wd

=-

9

WC + ___ = J . (1 - 9)

1-y)

(6-74)

(where V an apparatus and substance-specific variable, is similar to the terminal velocity of an individual droplet with phase

420

6 Extraction

velocity w, = 0 and small holdups of the dispersed phase (p-0). However, V is smaller than the real terminal velocity of single fluid particles in the continuous phase, and does not correctly describe the real physical behavior. Therefore, newer concepts based on droplet swarm relationships are used to calculate the relative velocity with p -+ 0, the steady-state velocity of a single droplet (see, for example, [6.29]). Phase throughput and phase velocities wdr w,,respectively, reach a maximum at the flooding point. The flooding point is the hydrodynamic loading limit of the column, exceeding its limit disturbs the phase flow. Either the dispersed phase is then carried out by the continuous phase, or the dispersed phase coalesces to streaks or drops. The column is now partially or total blocked, and phase inversion from the disperse phase to continuous phase occurs. The maximum values wdF and wcF, according to THORNTON [6.28], are obtained from Eq. (6-74) by differentiation of the phase velocities with respect to holdup

lid,

Following from Eq. (6-74) with the maximum value of the holdup pF in Eq. (6-78), (6-80)

and the loading limit of the column is

<

Usually, practical design correlations for wdFand wCFor the characteristic velocity V are directly derived from experimental data. For example, in an Rotating Disc Contactor (RDC), in the operating range of interest, the velocity limit of the dispersed phase wdF at the flooding point is according to STROBEL and SALZER [6.30]

.(e,

“)-“7

. (:)-*.27

.

giving Wdfi

= 2 * P . p i . (1

- VF)

(6-76)

and W& =

where

v ’ (1 - 2

’

9 ~ ’ )(1 - fJ?F)2

(6-77)

By elimination of V , the maximum holdup at the flooding point follows

(h2+ 8 * d)’.’ - 3 . A VF=

4.(1 - d )

with the phase ratio

(6-78)

d,, N,, Z, diameter of stirring disc, number of stirred cells, height of stirring cell n, g number of revolutions of the stirrer, gravitational constant e, q, 0 density, dynamic viscosity, surface tension A phase ratio For a defined system treated in a certain column under defined operating conditions, the characteristic velocity V and the

6.2 Liquid-Liquid Extraction

mean diameter of the droplets ds (Sauter diameter) of a droplet swarm reach a constant value. Therefore, ij is generally correlated as a function of ds which is, for example, for an RDC [6.24] 0.14 (g *

v=

a

. rl:.46

@2:7

. d,’.I9

(6-83)

where A@ is the phase density difference. A safe distance from the flooding point is given by considering the following relationships (c, column contraction factor, cross-section ratio) ~

Additional correlations for W d F , wcF, v, and d, for different extractor designs are found in special design literature of manufacturers and [6.18, 6.191. Further design considerations are found in Chapter 6.2.4. Loading diagrams to determine the loading limits for countercurrent extraction columns are presented by PILHOFER [6.31]. The holdup ~1 of the dispersed phase, fixed by size, number, and velocity of the droplets in a droplet swarm, determines the interfacial surface area referred to the volume A

-

V ~

v

wd *

= 0.16

c,

for i2 1.0

wd

0.476

2+A

“c,

for A

< 1.0

-

(6-87)

where Re, is the Reynolds number of the stirrer (Eq. (6-67)) and we is the Weber number

We =

d j .n2 CT

e,

(6-89)

6.2.3.5 Countercurrent Extraction with Extract Reflux

For the operating range of interest of the RDC (6 lo3 < Re, < 7 . lo4) d, is, according to [6.19] 0.17 . We-’.2.

6-V1 or a,=- 6 . 6 . 9 dS d,

(6-85)

(6-86)

dr

-a,=-

Formulae to calculate the mean holdup of different extraction columns are given in [6.18].

The Sauter diameter d, (d32)is the characteristic diameter of the droplet spectrum at steady state (ni number of droplets, d,; droplet diameter).

dS = -

-

(6-84)

and -=--

421

(6-88)

In countercurrent extraction, the concentration of the key component in the extract Em leaving the extractor may be increased by refluxing of treated extract phase to the column. Figure 6-24 shows schematically the principle of countercurrent extraction with extract reflux. The feed F is charged at the top of the raffinate stripping section of the column; the feed location is adjusted according to the concentration profile of the column. Fresh or regenerated solvent is used in the raffinate stripping section, and subsequently in the extract enrichment section of the column, an upflowing extract phase EE and downflowing raffinate phase A,, are formed. In the top section of the column the extract phase is further enriched with the key component and then leaves the column as gm.After separation of the solvent Lo in an extract separation unit, a key component-rich solution phase io remains,

422

P

1

6 Extraction

i,

Em = L o -k P

su ‘0

-k

(6-91)

R,

If (6-92)

L,+P=U

the net flow out of the top of the extractor it follows

Em=

u +R ,

(6-93)

In the triangular diagram in Fig. 6-25 the state point U of 0 lies on the connecting line between the state points E, and Lo of the flows Em and Lo and according to Eq. (6-90) also on the line A mass balance over the extract enrichment section of the column (Fig. 6-24) at any cross section gives

35

m.

?S

---t

Fig. 6-24. Countercurrent extraction with extract reflux. SU Extract separation unit EC Extraction column RS Raffinate stripping section, below feed point ES Extract enriching section, above feed point

Therefore, the state points of the extract and raffinate phases E and R in a common column cross section lie on the binodal curve and connecting line E,R,U (lines through the “pol” U) (Fig. 6-25). The state points of the phases leaving a theoretical extraction stage (Sm, R , ; E l , R,; etc.) lie on the binodal curve and common tie lines. The number NE of theoretical stages of the extract enrichment section is then obtained by a stepwise construction of pol lines and tie lines. With the Lever rule for the phasedmixtures, from Eqs. (6-92) and (6-93) (6-95)

which is divided into the product flux P and the column reflux R,. A mass balance over the extract separation unit gives

The extraction reflux ratio (“external reflux ratio”) as the ratio of the extract reflux R, and extract product flux P is then

Eu = L, + E,

\IE=-=7.-

and including the reflux R ,

(6-90)

Ro

P

.

E, Lo

~

_

_

_

_

_

_

E,U/URo - _ PL, PU*E,U - __._

uL,/pu PE,UR,.UL,

(6-96)

_

6.2 Liquid-Liquid Extraction

Since Pu =

m, it follows that

Replacing dE- R.E= E - R and E R - R R = (6-97)

Therefore, the location of the state point U is clearly defined by the desired extract reflux ratio vE. The required number of theoretical separation stages NR in the raffinate stripping section of the column is determined in a similar way to that for the simple countercurrent flow of extract and raffinate phases. A mass balance over the mixture feed section gives (Fig. 6-24)

ER t R E + F

=EE

+ R,

(6-98)

or

W+F=U

=

423

U (Eq. (6-94)) (6-100)

This implies that the state points W, F and U of the flow rates % and U must lie on one line (Fig. 6-25). W lies at the intersection of this line with the extension of the line R,L, beyond La (see Chapter 6.2.3.4). The number of theoretical stages N R of the raffinate stripping section are determined by considering that the state points of the equilibrium phases lie on common tie lines and the binodal curve. The state points of the phases in a common cross section must lie on the common line connected with the pole W (“pole line”). For example, N, = NE + NR corresponds to the number of the state points of the raffinate phase R,, ..., R, on the binodal curve. ~

s

T

Fig. 6-25. Graphical determination of number of theoretical stages, countercurrent extraction column with reflux.

424

6 Extraction

La = Lo f P 4- R , - F

(6-101)

6.2.4 Design Forms of Extraction

Apparatus 6.2.3.6 Countercurrent Distribution

If a mixture, consisting of two components S, and S,, is treated with two insoluble solvents L, and L, so that S, dissolves L, and S, dissolves L,, countercurrent distribution has to be applied (Fig. 6-26). The feed is charged in the middle of the extraction column and the solvents are fed at the top and the bottom of the column. During the phase contact, L , and L, are dissolved in and according to the solubility. In an ideal case of high selectivity, the feed may be totally separated. Two columns are often

s,

s,

:

i2

il+ 5,

The extraction process begins immediately after dispersion of one phase into the other and ends when the dispersed phase forms a continuous phase after phase separation. To exchange as much of the key component as possible from the feed to the solvent during the extraction process, the extraction unit must fulfil the following requirements: 0

0

7r 0

I I I I I I I I

s,

5,+ v L

"

0

0

tI I I I I

I I

I 1r

i2+ 5,

,, Ll

Fig. 6-26. Countercurrent distribution.

A large interfacial area for mass transfer, fast droplet formation with a narrow size distribution, and fine distribution of these droplets in the continuous phase to provide a large interfacial area for mass transfer Large flow velocities of the phases and hence, a high phase turbulence and suitably high mass transfer coefficients Large mass transfer driving forces along the total length of the extractor, through optimum phase flow (guided phase flow) with minimum backmixing Sufficient time for phase contact Good phase separation

Technical extractors are usually constructed with a maximum of 10- 15 theoretical separation stages in a single apparatus. Compared with rectification, this reduced number of stages is due to the small stage efficiency (exchange ratio, enrichment ratio, HETS, etc.) caused by the more difficult mass transfer between liquid phases. Mass transfer mainly occurs by diffusion with low key component concentrations and is favored by large, turbulent, constantly renewed interfacial areas. In different extrac-

6.2 Liquid-Liquid Extraction

tor designs, the renewal of the interface area can be achieved, for example, using internals, energy dissipation by pulsation, stirring systems, etc. The main reason for the low separation effect is the more or less pronounced backmixing of both phases, which increases with increasing diameter of the extractor. Figure 6-27 gives a short overview of extraction processes and extraction apparatus. Common extraction apparatus are classified in Table 6-4. In the following, essential features of important designs of extraction apparatus are described in tabular form. Design considerations are included. For further detailed treatment and special design considerations of these extractors, literature is cited.

425

6.2.4.1 Mixer-Settler, Mixer-Settler Cascade The mixer-settler (Fig. 6-28) consists of a mixing chamber (mixed vessel, pump, mixing nozzle, static mixer, etc.) and an adjoining separator, which is separated by a slit plate from the mixing chamber. In the mixing chamber, the feed and solvent are intensively mixed and remain in contact for the duration of the mass transfer of the key component. The mixing device has to be operated such that an optimum droplet size is obtained, i. e., a compromise between the small droplets favored for mass transfer and large droplets for small phase separation times. An optimum stirrer revolution speed is found from the minimum of the curve for

Extraction

I Solid-liquid

I Gas-liquid

I Liquid-liquid I

I

I

Single stage

Multistage

I

I I Mixer settler

1 (separator)

I 1

Cocurrent Centrifuge extractor I

I

Mixerlsettler Separator

l l

Countercurrent

I

I

I

7

Without energy input

With energy input

d With pulsation

With mixer

I

Individual units in series

Chamber separator

Podbielnak

Spray Scheibel column

Box type

Disc separator

Quadronic

Packed column

Tower type

Kuhni

Sieve tray RDC column ARD and others

I

Packed column

Sieve tray column Prochazkacolumn Karr-column

Fig. 6-27. Extraction operations and extractors. Representation according to Brandt et al. t6.31.

426

6 Extraction

Table 6-4. Classification of selected commercial extraction apparatus with countercurrent phase

flow.

Countercurrent phase flow is caused by

Gravity

Mixing and distribution of the dispersed phase in the continuous phase is achieved by Extraction apparatus with continuous phase contact

Gravity

-

Stagewise contact with controlled coalescence redispersion cycles

Centrifugal force Pulsation

Mechanical agitator

Centrifugal force

Spray column Pulsed packed Packed column column

Rotating disc contactor, Oldshue-Rushton column, GraesserContactor, Kiihni column

PodbielniakExtractor, Lurgi-WestfaliaExtractor, De LavalExtractor

Tray column

Scheibel column, ARD-Extractor, Leisibach column, MixerSettler cascade

Pulsed sieve column, Pulsed MixerSettler-cascade, Extraction tower with controlled cycle

Table 6-4a. Classification of extraction columns. 0

0

0

0

Column (packed column, sieve tray column, packing column) Pulsed column (packed column, sieve tray column, Karr column) Agitated column (rotating disc contactor (RDC), asymetric rotating disc contactor (ARD), QVF agitated cell extractor, Kiihni Extractor etc.) Mixer-settler tower

the sum of the residence times in the mixer and the phase separator. The settler is usually a horizontal vessel filled with a n easily wettable coalescer (filling material o r wire packing at the entry, slanting mounted sheet metal packets at the front of the settler section, steel wool, etc.) to increase the separation efficiency. I n process practice, u p to

5 mixer-settlers are operated in a cascade with countercurrent flow. A box o r tower form avoids the expensive single stage set up a n d allows a larger number of stages. However, the separation effect is then reduced. Table 6-5 presents some important con4derations for the design a n d operation of mixer-settlers a n d describes the advantages a n d disadvantages.

6.2.4.2 Countercurrent Columns with and without Energy Supply I n countercurrent flow extraction columns, the heavy phase is charged at the top of the column and t h e light phase at the bottom of the column. Vertical countercurrent flow of the phases is caused by gravity a n d density difference. If t h e light phase is the dispersed phase, the phase separation level has to b e high. The reverse is true in the case of the heavy phase as the dispersed phase.

6.2 Liquid-Liquid Extraction

427

Table 6-5. Hints for design and operation of mixer-settlers [6.23, 6.31-6.331.

Loading range ca.

100 m3/h

Stage efficiency ca. 90-95 Vo (single arrangement) Design based on experiments in a pilot-scale unit under conditions as close as possible to operating conditions. Thus, scale up is no problem Design hints for horizontal separators [6.23]: 0

Minimum separation volume Vmin

v

tdyn 0

%in =

‘’

(6-102)

tdyn

Volumetric phase flow rate Dynamic separation time

Minimum separation length Lmin

Lmjn= vm/min/AQ = v ’tdyn/!Q

0

Separation cross-sectional area A Q (6-104)

A p = V/W,II 0

Allowable velocity wall = (0.5

0

1

(6-103)

. . . 0.9).

WF

(6-105)

Flooding point wF

w F = 0.01 m/s (approximate value)

(6-106)

Mixer-settler advantages: 0 0

0 0 0 0 0

Large loading range, good stage efficiency Suitable for extreme phase ratios Adjustment of optimum droplet size or optimum residence time by an appropriate design of the mixing chamber Treatment of even solid-loaded phase fluxes Horizontal cascade expandable Simple start-up and shut-down procedure Simple design and, therefore, simple maintenance

Mixer-settler disadvantages: 0 0 0

I

High energy consumption and high process control cost Large hold up and, therefore, high solvent cost Horizontal cascade require large area The phase with the lower volumetric flow disintegrates usually to droplets, a choice of which phase has to be dispersed is only possible in a small range of phase ratios

428

6 Extraction Aeration

a)

t

LP

LP

HP Sludge outlet HP +-! Interface sludge outlet

-

Mixing zone

Settling zone

Fig. 6-28. Design and operating mode of mixer-settler types. Representation according to BRANDT,REISSINGER, S C H R ~ E(6.311. R LP Light phase HP Heavy phase a) Mixer-settler b) Box-type mixer-settler cascade c) Principle of a Lurgi-tower extractor

6.2 Liquid-Liquid Extraction Cl

HP

429

LP

Mixer

valve

t

LP

t

HP

Countercurrent Flow Columns without Energy Input

Countercurrent flow columns without energy input operate without external influence on the liquid flow and droplet distribution. To generate droplets in the column, only the potential energy of the liquid system is converted and used. An increase in the throughput leads to smaller droplets; a large fraction of the dispersed phase (holdup) should be achieved under reliable operating conditions. In such columns the separation is relatively poor due to the unfavorable droplet size distribution and the unsatisfactory interface renewal. The technique is only applied if the solution to be extracted requires a small number of separation stages. Spray columns, packed columns, sieve tray columns, and perforated plate columns are shown in Fig. 6-29.

Fig. 6-28c

In a spray column with internals (Fig. 6-29a) [6.34, 6.351 droplets are formed mainly at the feed location by means of annular sprinklers. Droplets now move under the influence of gravity and buoyancy forces through the continuous phase. The column throughput depends considerably on the phase density difference and the viscosity of the phases. Considerable axial mixing is disadvantageous. This increases with increasing column diameter/column height ratio. Saddles or rings with perforated walls (Berl saddles, Pall rings, etc.) are used as the filling material in packed columns (Fig. 6-29c) with individual packing sizes ca. 15-25 mm for random packing. In this case, or for regular packing with, for example, Bialecki rings, the flow control through the packing is favorably influenced. The chosen column diameterkharacteristic packing dimension ratio, for the case of

430

6 Extraction

Fig. 6-29. Countercurrent extraction columns

[6.98, 6.991. a) Spray column b) Perforated plate column c) Column with irregular packings d) Column with regular packings, Sulzer-Mixer packing SMV * LP Light phase FM Filling material IF Interface LIC Level indicator and control HP Heavy phase ST Sieve tray

* Representation according to data of Sulzer, Winterthur.

random packing, should be larger than 8. The size, form, and arrangement of the packing determine or even control the dispersion and mixing behavior, the velocity distribution, and the phase residence time for any given system and operating load. Good distribution of the disperse phase is essential at the entry to inhibit the formation of streaks. The continuous phase has to show good wetting behavior with the packing. If the disperse phase wets the packing, the coalescence increases and the interfacial area decreases. In extraction columns with regular packing, the separation effect is increased by

Ic

HP

6.2 Liquid-Liquid Extraction

blowing inert gas in with the dispersed solvent. Additional energy is supplied to the column by the gas which causes a better droplet dispersion and, therefore, an increase in the dispersed phase fraction and with it an increase in the mass transfer area [6.103]. Simple packed columns are only used for simple separation and for systems with relatively large density differences (A@> 150 kg/m3) and a surface tension in the range of 0.005 N/m < o < 0.015 N/m. In columns with regular packing (Fig. 6-29 d), packing elements with a regular structure (for example, corrugated fins) such as the mixer packing SMV of Sulzer are employed. The packing is distinguished by pronounced cross-mixing and low axial backmixing, thus allowing high throughputs with small density differences and surface tension, and making good partial loading behavior of the column possible. Literature for the design of columns with random and regular packing is found in [6.36-6.381. Table 6-6 gives some hints for the rough estimation of extraction column dimensions with random or regular filling material or packing, and estimated data for the loading range and separation efficiency. In perforated plate columns (Fig. 6-29b) [6.42], perforated plates or sieve trays are employed to distribute the disperse phase into the continuous phase. For example, if the light phase has to be dispersed, it is fed at the bottom of the column. The light phase is redispersed at each tray and coalesces while flowing upward through the column. The heavy phase is guided in a cross-flow manner to the light phase at each tray by means of downcomers guiding the flow from tray to tray. On the trays, mass transfer zones develop in which the light phase passes through the heavy phase in the form of a droplet swarm. A clear layer of coalescing light phase is found above the mixing zone, which is extended to the next tray. (In systems where the dispersed phase

431

slowly coalesces, an intermediate zone consisting of stacked up droplets develops between the mixing zone and coalescing zone.) The buoyancy force, generated by the volume of the dispersed phase between the sieve tray bottom edge and the exit level of the downcomer (see Fig. 6-33), provides the required energy to disperse the light phase in the area of the sieve tray. With small throughputs of the disperse phase periodic droplet formation occurs by droplets individually detaching from each hole. With increasing disperse phase throughputs, droplets are formed by jet disintegration; the disperse phase leaves the hole as a jet, and at the end of the jet, single droplets are formed. Since the dispersed light phase coalesces below each tray, axial backmixing in the continuous phase is limited to the space between two sieve trays. The operating range of a sieve tray column is determined by the hole size opening ratio, downcomer length and the phase throughput. In systems with a low surface tension, the hole size of the sieve tray with downcomer should be smaller than 2 mm to prevent the continuous phase percolating through. For the treatment of such systems, dual flow trays without downcomers are more suitable. On a dual flow tray, the disperse and continuous phases flow in turn through the base plate holes and between the trays, producing strongly circulating convection cells. Table 6-7 gives some advice for the estimation of the sieve tray column dimensions and operation methods.

432

6 Extraction

Table 6-6. Hints for approximate design of extraction columns with selected filling material and packing’). Specific load: ca. 20 m3/(m2 . h) Packed column with irregular packings in the range of maximum separation efficiency ca. 100 m3/(m2 * h) Sulzer-SMV column (depending on substance system) 0

Filling material dimension (diameter) dp: (6-107) (at dp < dpkdroplets get caught between the filling material and coalesce with the following droplets [6.39])

0

Droplet size range (packed column [6.40]): 1<

0

d:[email protected] <9 a

(6-108)

Characteristic droplet velocity v, given in m/s [6.38]: (6-109)

0

Phase flow velocity W d F , wCF,in m3/(m2. s) and holdup pF at the flooding point (see Eqs. (6-72) - (6-81)) [6.38] WdF = 2

’

v ‘ & ’ P,?(l - V F )

v

’

&

WcF =

(A2

” 0

(1 - 2 ’ V F ) * (1 - pF)2

+ 8 . A)’.’

4(1 -A)

- 31

wd C0’E.V

for w,< 0.65 ’ wcF

(6-112)

(6-113)

(a, increases with decreasing droplet size and increasing phase ratio A, increasing droplet size increases the flooding point) Specific mass transfer area (interface) a, in m2/m3:

6 . p . ~

a, = -

0

(6-111)

Dispersed phase holdup a, < 0.25 m3/m3 [6.38]: a,=-

0

=

’

(6-110)

ds Height of a transfer unit HTU (without pulsation), given in m [6.38] Mass transfer direction d-c:

(6-114)

(6-115) Mass transfer direction c-td: (6-116)

’)

For the final design experiments are required under conditions as close as possible to operating conditions as well as scale-up experience.

6.2 Liquid-Liquid Extraction

433

Table 6-6. (continued) Nomenclature for Eqs. (6-107-6-116)

4

Sauter diameter of droplets Voidage 1 Phase ratio, volumetric m3/m3 @ , A @ Density, density difference of phases in contact, kg/m3 0 Surface or interfacial tension D Diffusion coefficient, m2/s Slope of equilibrium curve, kg/kg rn C, C,, C3 Specific packing constants [6.38] Index: c continuous phase, d dispersed phase V individual droplet velocity &

Packing

&

C

co

c 3

Bialecki ring 25 mm, metal, regular

0.928

0.67

0.465

267

Corrugated sheet metal pack B1-300, Montz

0.93

0.60

0.415

445

Pall ring 25 mm, metal, random

0.94

0.40

0.336

346

Hiflow ring 0.83 35 mm, ceramic, random

0.57

0.47

370

0

Pulse effect: higher load, approximately separation efficiency doubled.

0

Operating diagram, separation data, see Figs. 6-30-6-32. Fig. 6-30. Separation efficiency of pulsed and unpulsed columns (experimental column 0 72 mm. Representation according to BRANMet al. [6.31]).

-

1

0

fit

a .f

FP

6+

5

10

i', + i', [ m 3 / l r n z ~ h )l 15

20

25

= No. of theoretical stagedm = 6 * 96 = 576 mm/min = Flooding point

= Cross-sectional load

30

35

LO

System: Water-Acetone-Toluene Pulsed column: -0VA-Interpack 15 x 15 mm -0- VA-Pall rings 15 x 15 mm -ACeramic saddle 15 x 15 mm Unpulsed column: - - 0 - - VA-Interpack 15 x 15 mm --m-- VA-Pall ring, stainless steel 15 x 15 mm --A- - Ceramic saddle 15 x 15 mm (continued next page)

434

6 Extraction

Table 6-6. (continued)

Fig. 6-31. Operating diagram of an extraction column with SMV packing. Representation according to STREIFF [6.41].

80-

1 5

01 0

a1

2-

1 . HTU.3,

d-c

1 10

II

20

wC

T-T

c-d

1-

I

15

I

-

25

[mm/s]

\ FP 0-0

h c F P

C-d

Pulsed

(1200mm/minl

‘<x4-T-*Fp 0

0

0

20

LO

f.,+ r‘, [m’/lm’.h)]

6-32. Separation efficiency data of an extraction column with SMV-packing. [6.41]. Representation according to STREIFF

~ _ _ Fig.

~-

I

1 1

System: Tohenelwater q~ Disperse phase holdup wc True phase velocity of continuous phase I w d True phase velocity of 30 dispersed phase

-

60

System: a) Toluene/Acetone/Water b) n-Butanol/Butanedioic acid/Water FP Flooding point , I= 1; Organic phase dispersed

6.2 Liquid-Liquid Extraction

435

Table 6-7. Hints for the estimation of sieve tray column dimensions, design and operation [6.19], [6.42].

Principle (Fig. 6-33). Representation according to BLASSet al. [6.19] and PILHOFER et al. [6.42].

0

b!

LP

HP

LP

HP

Restrictor layer

LP ~

1

1

LP

HP

Fig. 6-33. a) Sieve tray with downcomer

b) Dual-flow tray

DC LP, HP DP, LP CL CC

Downcomer for continuous phase Light and heavy phase Dense and loose packing Coalesced layer Circulating convection cell

0

Specific load: up to ca. 70 m3/(m2 h)

0

Suitable for systems: A @ > ca. 100 kg/m3;

0

Tray geometry:

17> ca.

0.005 N/m

Hole diameter db 2 mm Opening ratio psA = 5 % (sieve tray with downcomer) pDA = 10. psA (Dual-flow tray) Column diameter d : up to ca. 4 m Tray separation: ca. 120-500 mm (continued next page)

436

6 Extraction

Table 6-7. (continued) 0

Sauter diameter d, (Fig. 6-34) Calculation of d,, for example see [6.19]

Fig. 6-34. Sauter diameter as a function of hole diameter. Representation according to BLASSet al. 16.191 and BENDERet al. t6.431.

Process data: Single holes of diameter 1.5 mm.

System: ToIuene/Water and ButanoVWater. d, Mean droplet size

utanol/water

0

0

LO 60 80 Mean velocity of disperse phase In holes Icrn/sl

-

20

Restrictor layer height

100

zs [6.42]

(6-117)

APb

Sieve tray pressure drop (6-118) where Reb =

APA

wh

@A

w1,

'Id

(6-119)

~2

(6-120)

Mean density of two phase mixture in the downcomer region

ern = p .@d+ ZA

dh ' @d

Downcomer pressure drop ApA = 2.41 . e,.

ern

*

- p) ' e c

(6-121)

Continuous phase density in downcomer Downcomer length (to be fixed, thus the pressure drop of the tray determines the restrictor layer thickness and not the pressure drop of the downcomer) Dispersed phase velocity in the hole of diameter db

6.2 Liquid-Liquid Extraction

437

Table 6-7. (continued) 0

Loading point, operating range [6.42] (Fig. 6-35) Fig. 6-35. Volumetric flow of disperse and continuous phase at the loading limit and operating range of a sieve tray. Representation according to REISSINGER et al. [6.23] and PILHOFER et al. [6.42].

Process data: Downcomer area 0.02 m2, hole area 0.025 m2, ratio of downcomer length over hole size 250, dimensionless velocity

w.*=

0.1 0.1

0

0.2

-

0.L w;

0.6 0.81

2

@IQ

( A @ .V , . g / ~ , ) ' ' ~

(6-122)

w,* Dimensionless continuous phase velocity W $ Dimensionless dispersed phase velocity

Separation efficiency [6.42] Approximate values of the enrichment ratio : 0.2-0.5

Advantages and disadvantages of sieve tray columns 0

- Simple design with simple flow path

- No axial backmixing between stages

- Droplet size adjustable by hole diameter, narrow droplet size

I

distribution

- At constant fluid mechanics scale-up possible - Large cross-sectional load 0

I

Advantages:

Disadvantages: - Sensitive to fouling and incrustation - Low energy density for phase dispersion - Large section of column height for clear restrictor layer but not usable for mass transfer - Only suitable for systems where A @ > 100 kg/m3 and cr > 0.005 N/m

438

6 Extraction

Countercurrent Columns with Energy Input In countercurrent columns with energy input, pulsation or agitation causes an improvement in the renewal of the droplets and the surface, and hence an improvement in the separation effect, particularly with systems having a large surface tension. In pulsed, packed and sieve tray columns, pulsations of short amplitude are induced in the phase mixture by a pulsator (piston pump with a blind flange on the suction side, without valves, (pulsation piston) or flexible, mechanical or hydraulic, swinging, metal or synthetic, skins) including an intermediate gas bolster, if necessary. With a constant stroke of a = 6- 10 mm, the pulsation intensity a f is varied by varying the frequency, f I 150 min-'. Pulsation causes homogeneity in the two phase flow across the column cross-section. This increases the mass transfer area through the formation of smaller droplets, and widens the spectrum of systems which may be handled. The maximum phase throughput is not significantly increased by pulsation but the separation effect is almost doubled. In pulsed packed columns [6.19, 6.431, the loadability decreases with increasing pulsation frequency, but generally increases with larger dimension of the filling material and an increasing void fraction. The type of mass transfer, continuous -+ disperse or disperse continuous, generally influences the droplet motion and separation efficiency :

-

+

0

0

0

Increases in the key component concentration reduces the surface tension of the extraction system Mass transfer in the direction disperse continuous promotes droplet coalescence Mass transfer in the direction continuous disperse hinders droplet coalescence, and therefore supports the formation of small droplets +

-+

The above influences of the mass transfer direction are particularly pronounced for systems having a high surface tension. In these systems, mass transfer from disperse continuous allows a considerably larger loading limit than reverse direction. The separation efficiency increases with increasing pulse frequency and smaller filling material dimensions. Mass transfer from continuous disperse is more favorable here. Table 6-8 gives some information on aspects of loading behavior and separation efficiency of pulsed packed columns. There are two different types of energy input in sieve tray or perforated tray columns (Fig. 6-39): in a pulsed sieve tray exfractor (PSE)the liquid column is pulsed; in a swing tray extractor (STE) fixed sieve trays are mounted on a swinging axis which is driven by an infinitely variable, directly coupled geared motor, including crank shaft and connecting rod. The top and bottom section are expanded to allow the phases to rest and then separate. With low pulse intensities, the disperse phase is collected at the base plate and gradually conveyed through the sieve tray holes. With upward motion the heavy phase is dispersed, with downward motion the light phase is dispersed (mixer-settler range). The droplet size remains virtually constant and the holdup (large) depends on the pulsation. With increasing pulse intensity, the column begins to operate in the dispersion area; the build up of droplets vanishes. Droplets become smaller and remain in the space between the trays. The PSE operates between the mixer-settler and the dispersion range with frequencies between 60 and 150 min-', and amplitudes below 10 mm. Maximum throughput is then ca. 60 m3/(m2. h) with minimum holdup and a good separation efficiency. The best operating point for the PSE is with a pulsation frequency of 10-20 strokedmin above the pulsation frequency for the maximum flooding ooint r6.441 (Fie. 6-37 in +

+

Y

.

L

A

\

-

439

6.2 Liquid-Liquid Extraction

Table 6-8. Hints for design and operation of pulsed packed columns. Optimum cross-sectional load: ca. 20 m3/(m2. h) 0

Density difference: Ae > 80 kg/m3 Loading point, operation characteristic (Fig. 6-36)

1

Fig. 6-36. Loading point of a pulsed packed column, system Toluene/Acetone/Water. Representation according to PILHOFER and SCHROTER [6.46].

90

fd + Pc

80

[rn3/(rnZ hl]

60

Vd +

50

f

LO

a

30

- a

< Total cross-sectional load Pulsation frequency

e/pd

-- a

Stroke length = 0.55 =8mm =6mm

Packing

o ! 0

1

1

50

100 f [l/min]

-

I

1

200

150

0

x

A

+ V

0

E

0

E

(Vo) Mass RI transfer

Pall-15 stainless steel 93 Raschig-25 ceramic 73 Pall-25 stainless steel 94 Berl-15 ceramic 61 Interpack-15 stainless steel 93 Pall-15 stainless steel 93 Sulzer-SMV >95

t

+

+

16 I6

+

Voidage

Separation efficiency, residence time 0 (Eq. (6-123) 8

Fig. 6-37. Separation efficiency of pulsed packed columns. Representation according to BRANDT et al. [6.31].

7

6

‘ml

nt

Number of theoretical stages per m column length a . f Pulse intensity

5 4

3 0

500

700

-

900 1100 a . f [rnm/min]

1300

Example: Pulsed packed column 0 72 mm with 15 x 15 mm Interpack packin Sample mixture: Water-Acetone-Butyl acetate Load: $ + pd -A- 15 m3/(m2. h), -o- 20 m3/(m2 h) -x-25 m3/(m2. h) - ~ - 3 0 m3/(m2 h) FP

=

Flooding point

440

6 Extraction

Table 6-8. (continued) a)

1

Fig. 6-38. Residence time 0. Representation according to BENDERet al. [6.43].

0.08

a) System: Toluene/Water/Acetone b) System: Ethyl hexanol/Water/ Acetone

0.06 0 [hl 0.04

0.02

x

____

Raschig 25 ceramic Pall 15 stainless steel Pall 25 stainless steel Brand et al. [6.31]

0 f

Residence time Pulsation frequency

00

0

0

50 f

-

100 [min-ll

150

0

50

-

100

~[min-']

150

V

wire gauze, grids, packing, etc.) limits and subdivides the individual mixing cells into single chambers in the column. Phase dispersion and mixing occurs in HE TS (6-123) the mixing cells where transfer of the key 0=vd -k vc component takes place. The characteristic droplet diameter is barely influenced by a where 0, the residence time of a theoretical low number of mixer revolutions n, (of little separation stage, is a minimum and, there- interest in most technical applications) ; mixing has little influence on the interfacial fore, is the required column volume. The aperture ratio of the PSE has to be area, the column operates as a spray coladjusted to the substance system accord- umn. The dispersed phase has a low holdup ingly. Small surface tension requires a large and is unevenly distributed over the column aperture ratio and hence larger hole diame- volume. By exceeding the lower revolution ters. limit ng, the agitated column reaches its With an increasing aperture ratio the operating range. The droplet size decreases loading limit increases but the separation with an increasing number of revolutions efficiency decreases. and the interfacial area increases. Due to Table 6-9 gives some information on de- the smaller droplet size, however, the dropsign, dimension estimates and operation of lets rise more slowly and, therefore, an increase in the column diameter is required PSE and STE. With mechanically agitated coun tercur- which unfortunately leads to axial backrent extractors, energy is induced by rotat- mixing (unequal radial distribution is coning internals. Agitators are usually mounted siderably avoided by mixing in axial direcon a central column shaft at certain dis- tion). An increased number of revolutions tances from each other (double blade mixer, also necessitates a larger energy requiredisc mixer, blade mixer, turbine mixer, ment. However, with smaller column diamhelical screw mixers, etc.). Installation of eters, the power consumption of the mixer stators (ring discs, punched metal sheets, shaft is only partially used for the break-up Table 6-8). From the separation efficiency HETS and the throughput r'd +

vc

441

6.2 Liquid-Liquid Extraction

of droplets, due to increasing losses through wall friction. Phase separation occurs in the resting area of the stator elements where easily wettable internals promote coalescence. The column separation efficiency increases with smaller aperture ratio and larger contraction of the column cross-section in the area of the stator element, but also reduces the phase throughput. Commonly used agitated columns in extraction processes are shown in Table 6-10. Figures 6-41 to 6-43, provide information on operating ranges, loading limits, and separation efficiencies of selected agitated columns. Respective data are also given for certain substance systems with a certain mass transfer direction and apparatus geometry. Transformation to full-scale apparatus is only possible based on experimental scale-up knowledge (see, for example, Eq. (6-56)). Optimized dimensions of experimental extractors cannot usually be transferred to technical extractors, though the enlargement in an axial direction is generally smaller than in a radial direction. Agitated columns are characterized by the following advantages and disadvantages : 0

0

0

0

Large column diameters are possible and even allow large throughput of viscous mixtures, though with relatively small cross-sectional loadings Due to complicated design and, in some cases, poor construction, only limited heights are possible The droplet size is a function of the number of revolutions of the mixer, though a wide droplet size range results The mixer allows good cross-mixing; longitudinal mixing is essentially a function of the stator aperture ratio although it has a slight dependency on the column diameter

The torque Mz in a mixing cell is Mz = c w . n 2 . dr5 . ec

(6-126)

where c, is the drag coefficient, depending on the Reynolds number of the rotating disc (6-127)

(n revolution number of mixer, dr disc diameter, e, density and v, kinematic viscosity of the continuous phase). The required power Pz for each rotating disc is P, = 2.71. n .M,

(6-128)

Figure 6-44 shows the torque Mz as the Newton number, Ne MZ d,? . n 2

Ne =

- e,

(6-129)

as a function of the Reynolds number for a rotating disc contactor RDC. In a centrifugal extractor [6.20, 6.711 phase mixing, phase countercurrent flow, and phase separation occur in a centrifugal field. Individual centrifugal extractors can be distinguished by the type of phase control and the operating mode (chamber and disc separators as mixer-settler and as differential contactor, for example, Podbielniak and Quadronic extractors). Dispersion to single droplets occurs at the holes of the sieve inserts or at the overflow edges of the internals. In countercurrent extraction columns the settling time and separation area are calculated through the product of A Q . g (A@ density difference, g gravitational acceleration). For centrifugal extractors, g is replaced by the centrifugal acceleration b = r w2.b may be a multiple of g depending on the rotor diameter r (distance of the fluid element from the axis of rotation), the A

442

6 Extraction

6.2 Liquid-Liquid Extraction

443

I I I

7

1

I I I I

I I I

I

I I

I I

I I I

I

I

I

3

I

2

-?*

I

I I I

I

r

L

T

Fig. 6-39. Pulsed column with perforated plates and Karr column. Energy input by pulsation of liquid or induced by swinging sieve trays. a) Pulsed sieve tray extractor (PSE). Representation according to Stahl Apparateund Geratebau GmbH, Viernheim [6.45]. 1 Pulsation piston unit consists of 1.1 Pulsating piston 1.2 Operating piston 2 Stroke length control 3 Oil set, consists of 3.1 Oil tank and control valve 3.2 Main driving pump 4 Purge unit

b) Montz liquid-liquid extractor system KARR (swing-tray extractor). Representation according to Julius Montz GmbH, Hilden. 1 Column mass transfer zone 2 Separation chamber 3 Light phase storage 4 Heavy phase storage 5 Interface level control 6 Double feed pump 7 Movable column internals 8 Drive

444

6 Extraction

Table 6-9. Design, dimension estimates and application of PSE and STE [6.18, 6.19, 6.23, 6.44, 6.49-6.51, 6.1021. 0

Dimension, tray geometry, pulse intensity (approximate values) Column parameter

PSE

STE

Column diameter (m) Column height (m) Tray spacing (mm) Hole size (mm) Opening ratio (Yo) (relative open area) Amplitude (mm) Frequency (min-I) Loading (m3/(m2. h)) Separation efficiency n,('/m)

ca. 2.5 ca. 10 50- 150 2-6.5 5-25 (60)

ca. 1 ca. 10 50-150 6-15 30-60

5-10 60- 150 up to ca. 60 up to ca. 5

5-20 60-350 up to ca. 80 up to ca. 5

Pulsation frequency at flooding point maximumf,, minimum, Eq. (6-123) [6.48]), for PSE

f,

= c,

in min-] (residence time 0

vd . u + cj . u2 + (c4 + c, . u). In -

+ c,

(6-124)

v,

Amplitude 6 mm

Surface tension in N/m Opening ratio Constant

U

9s c, k

Constant

1

2

3

4

5

6

7

C

29.450 24.528

6.679 2.537

-0.1082 -0.0548

-2.067 -1.455

-0.426 3.247

0.1778

0.0437

k

Other flooding point correlations in [6.51] Operating characteristics, separation efficiency, residence time parameter of PSE [6.44] (Fig. 6-40) a1

a) Total cross-sectional load as a function of pulse intensity. * System: Toluene/Water/Acetone * Representation according to BERGER, LEUCKEL, WOLF[6.44].

+----'7"tToluene/water/acetone

20

lo-0

~

I

6.2 Liquid-Liquid Extraction Table 6-9. (continued)

b)

I

I

0

b) Height of a theoretical stage as a function of pulse intensity. *

Toluene/water/acetone

1

V x

3.0

HETS [rn]

0

+

I I1 I11 IV

3 3 4

6.5

23

40 40 60

* Representation by BERGER,LEUCKEL,

2 .o

WOLF[6.44].

1.o n-butanol/water/butanedioic

I

0

I

500

I

I

1000 a .f [rnm/min]

acid

-

15

c) Residence time as a function of pulse intensity. * __ Toluene/H,O/Acetone

-. - Ethylhexanol/H,O/Acetic acid _ _ _ n-Butanol/H,O/Butanedioic acid

*

O

-

0

500 1000 a.fImrn/rninl

Representation by BERGER,LEUCKEL, WOLF[6.44].

L 1500

Advantages and Disadvantages 0

- Large total cross-sectional load, good separation efficiency Over a wide range droplet size variable by pulse intensity - Suitable for systems with low interfacial tension (17 > 0.001 N/m) or small density differences (A@> 30 kg/m3) Disadvantages: - Sensitive to fouling and incrustation - At trays with low opening ratio narrow loading range

Advantages:

-

0

445

446

6 Extraction

Table 6-10. Common extraction columns with rotating internals. Design, principles, characteristic data and references [6.23, 6.31, 6.46, 6.521. Principles Characteristic data

Agitated column type Design Rotating Disc Contactor (RDC)'),

[6.53]In a rotating disc contactor phase mixing is caused by a disc mounted [6.56] on a rotor. Stator rings subdivide the column into individual compartments to reduce backmixing. Suitable for large throughputs as well as mixtures containing soil. Column diameter: up to 8 m Column height: up to 12 m Max. cross-sectional load: ca. 40 m3/(m2 . h) Separation efficiency: 0.5- 1 m-l Density difference: h e 2 5 0 kg/m3 Maximum throughput: 2000 m3/h

*)

Variable speed drive Light phase outlet

--Rotor

disc

, JHeavy

phase outlet

Asymmetric Rotating Disc Contactor (ARD)3, -\

Partition

Baffle

Troy

Shaft Disc

path

Assembly

Ref.

In an ARD-extractor the asym[6.57] metrical stator consists of trays and [6.58] baffles to divide the column into [6.59] extraction zones and linked transfer zones. On the rotor mixing discs are mounted. The extraction zone is limited by the stator partition and is separated by the trays into chambers. Phase transport and separation take place in the settling zone behind the partition. Compared to the RDC smaller throughput but better separation efficiency. Column diameter: up to 4 m Max. cross-sectional load: ca. 20 m3/(m2 . h) Max. throughput: 250 m3/h Separation efficiency: 1-3 theoretical stages/m Density difference: h e > 10 kg/m3.

6.2 Liquid-Liquid Extraction

447

Table 6-10. (continued) Agitated column type Design

Principles Characteristic data

Scheibel Column') Oldshue-Rushton Column')

In a Scheibel Column phase mixing [6.60] is caused by blade mixers or blade [6.61] mixers with mounted baffles. Wiremesh packing or filling material of one to three times the height of the mixing zone acts as the settling zone. Operation mode is like the mixer-settler principle. Separation efficiency ca. 3-5 theoretical stages/m.

Scheibel

H P Heavy Phase LP Light Phase

Kuhni Kuhni

W

a b c d e f

central shaft mixing blade baffle breake plate mixing zone separation zone

Oldshw-Rulhton

Ref.

The Oldshue-Rushton Column is [6.62] similar to the RDC. Instead of a disc rotor a turbine mixer is used. The individual chamber height is larger and is approximately the same as the column diameter. Mixing zones are separated by narrow passage holes. Separation efficiency: 1-3 theoretical stagedm with narrow stator rings and 0.8-1 theoretical stages with wide stator rings.

In Kuhni columns mixing zones [6.63] (with centrifugal mixers) are sepa[6.101] rated by perforated plates where the opening ratio (free cross-sectional area) may be adjusted to the desired operating conditions. Possible operating are mixer-settler mode and dispersion mode. Column diameter: up to 3 m Max. throughput: 350 m3/h Max. specific cross-sectional load : ca. 50 m3/(m2 . h). Separation efficiency: up to 10 theoretical stages/m, depending on the free cross section of the stator disc and operating conditions.

(continued next page)

448

6 Extraction

Table 6-10. (continued) Agitated column type Design

Principles Characteristic data

QVF Mixing-Cell Extractor (MCE)2)

QVF-MCE: Mixing zones (with blade [6.46] mixer, four elements) are separated by [6.64] partitions with specially designed openings (central circular with meandershaped weir to promote a channeling of both phases). Behind the weir a settling zone is formed to separate the phases. The heavy phase flows through the opening in the upper weir section into the circular area and passes downward. Analogously the light phase flows upward. At different points each phase passes through the opening area to increase the throughput and to decrease back-mixing. Column diameter: up to 1 m Opening ratio: 10, 20, 40% Max. cross-sectional load: ca. 15 m3/ (m2. h) depending on opening ratio Separation efficiency: cell efficiency 40-60%, i. e., 5-8 theoretical stages/m.

w

B Blade mixer with 1 elements P Partition W Meander shaped weir

EC Extraction Column STEINER)’) (System HARTLAND,

c

Basic design

Grid installation

Ref.

[6.65] In an EC (enhanced coalescence) [6.66] column mixing zones (blade mixer) are separated by grids made of thin metal sheet to support coalescence and axial flow. Due to special treatment of the grid surface the disperse phase coalesces at the grid openings and blocks part of the free area. Thus, an opening ratio adjusted to the respective load is achieved. The EC column is marked by a high loading flexibility and sufficient separation efficiency over the total loading range. Separation efficiency increases at underload operation. Max. cross-sectional load: ca. 80 m3/(m2 . h) Separation efficiency: ca. 2-6 theoretical stageslm.

6.2 Liquid-Liquid Extraction

449

Table 6-10. (Continued)

Agitated column type Design

Principles Characteristic data

SHE E~tractor’)~)

In an SHE column (self-stabilizing[6.66] high performance) mixing zones (with [6.67] paddle mixer) are separated by special installations. Rotationally symmetric installations in the phase separation zone force each phase to pass through channels which are tapered in the flow direction. In the case of overload the dispersed phase is stacked up in corresponding channels to the edge of the installations. If the overload still increases, the disperse phase flows through the channels reserved normally for the heavy phase. By simple reduction of the throughput or reduction of mixer revolutions a stable operating point is restored. SHE columns are marked by a high specific cross-sectional load. Max. cross-sectional load: ca. 100 m3/(m2 . h) Separation efficiency: ca. 3 theoretical stages/m, low load dependency.

M Mixer

C Conical

installation

Ref.

Controlled phase flow in conical installations

(continued next page)

450

6 Extraction

Table 6-10. (Continued) Agitated column type Design

Principles Characteristic data

Graesser Contactor')

[6.64] The Graesser contactor consists of a horizontal column divided into in- [6.68] dividual chambers by partions and a rotor with circulating dipping tubes. Phases to be contacted flow countercurrently through the column. Phase mixing occurs in the chambers. Via a gap between the circulating partitions and the column shell the mixture is passed from chamber to chamber. Due to gentle phase mixing the Graesser contactor is especially suitable for systems which form emulsions that are difficult to separate. This contactor is unqualified for systems with large density differences and large surface tension. The throughput is small and the length of a theoretical stage is approximately the drum radius. Column diameter: up to 1.8 m Max. throughput: ca. 25 m3/h Max. cross-sectional load: ca. 1-2 m3/(m2 h) Residence time: 3-15 min per theoretical stage. 1

') *)

3, 4,

5, ')

Representation Representation Representation Representation Representation Representation

according to according to according to according to according to according to

BRANDT, REISSINGER, S C H R ~ E[6.31]. R PILHOFER, SCHROTER [6.46]. Buss AG, Winterthur. M ~ G L[6.63]. I QVF Glastechnik GmbH, Wiesbaden. GAUBINGER, HUSUNG, MARR[6.67].

Ref.

I

22

20

-t 15

!

-1

I T

' I

7

f

N

E m

E

I

'rn

i I

10

\ I

\

I

\

-1 l

ips= 10 8%

I

cl

5

0.2-

4& . f i t-I

120

110

1

I

t

-

160 180 200 n [rnin-'I

-I-t

220 2LO

I

260

E Ln 0.1 kW

I

Fig. 6-41. Load ranges and efficiency of a QVFMCE and a Kiihni column. Representation according to PILHOFER,

-l$:--: I

-v-i'

.

-

8-

12

14 13

10

--

8

+/

0

t

I

!

~

'

I

~

I stator Load range cross of sections a QVF-MCE ps as afor function different of free the d l o.2 number of mixer revolutions -E m, Water/Acetone/Toluene A Water/Acetone/Butyl acetate 0.1 B Total load w I n Mixer revolutions I1 Separation efficiency of a QVF-MCE and Kiihni column as a function of the mixer revolutions n and the total load B. 0

m3/(rnZ.hl

..

~~

A

3

L..,

4

F

-

8-

12

10

I

;,/

i~---

~

8

I I

rn3/(rnZ.hl

452

6 Extraction

-

6 [rn3/(m2.h)]

0

0

10

20

30

LO

b [rn3/(m2.h)]

50

60

6.2 Liquid-Liquid Extraction

453

t

4 Fig. 6-42. Loading limits and efficiency of

SHE extractor and EC column. Representation according to data of QVF Glastechnik GmbH, Wiesbaden.

Ne

10

a) Loading ranges of an SHE-extractor -Toluene-Water _ _ _ _ Butyl acetate-Water b) Efficiency of a SHE extractor System: Toluene/Acetone/Water d-c Mass transfer direction, column diameter 100 mm c) Efficiency of an EC column System: Toluene/Acetone/Water Parameter: Revolutions per minute n, Theoretical stagedm & Total cross-sectional load

Fig. 6-44. Torque M, as the Newton number, Ne of one cell as a function of the Reynolds number Re for a rotating disc contactor RDC. Representation according to HUSUNG[6.69].

Ne Newton number Ne =

Re Reynolds number d

Mz

. n 2 . Q, Re = n . d;/v, d,'

Column diameter Z , Height of mixing cell n Revolutions d,. Diameter of rotor disc Q, Density, continuous phase v, Kinematic viscosity, continuous phase

8 !

-1

I

I e 2 -

1 0

I

10

-

20 30 nr [rnin-ll

LO

50

Fig. 6-43. Separation efficiency of a Graesser contactor. [6.70]. Representation according to STICHLMAIR System : Toluene/Acetone/Water Vd/< = 1.5 Tube diameter 100 mm Active length: 1 m Cell length: 0.025 m n, Number of theoretical stages per m n , Rotor revolutions

angular velocity o = 2 - x . n, and number of revolutions n. Therefore, centrifugal extractors allow treatment of systems with low phase density differences (A@2 20 kg/m3) and a tendency to form emulsions. At the periphery of the rotating extractor rotors filled with liquid, there is considerable pressure of 100 bar or more. Therefore, the extraction process in a centrifugal extractor is not isobaric, and this may affect the location of the distribution equilibrium I6.721. The total throughput and allowable pressure in the inlet and outlet for both liquid phases fix the operating range of the centrifugal extractor. For the design (Sauter diameter, droplet velocity, pressure profile, residence time distribution, flooding load, etc.) see [6.71, 6.73 -6.751. Extractor selection and operatine Darameters have to be determined exper" I

454

6 Extraction

Table 6-11. Selected centrifugal extractors. Design, principles and characteristic data [6.20, 6.31, 6.711. Extractor type Design

Principles Characteristic data

Lurgi-Westfalia Extractor (Luwesta-Extractor) [6.20]

Disc separators, divided into mixing chamber and separation chamber with conical discs. The liquid is distributed in thin liquid layers thus causing a short droplet path. Holes in the discs form rising channels for the upward flowing liquid. The liquid is introduced to the first stage via an inlet pipe and mixed with the extract phase withdrawn from the second stage by a centrifugal catcher. Phase separation and discharge of the final extract phase occur in the disc unit of the first stage. Mixing of the raffinate phase with fresh solvent and phase separation in the disc unit of the second stage, discharge of the raffinate phase. In countercurrent flow the extract phase is brought into contact with fresh feed in the first stage. Since in multistage apparatus the throughput decreases and fault liability decreases, singlestage centrifugal extractors are favored. If necessary single stage apparatus are connected to multistage units.

I--.-

Design')

7-

____-

-Heavy phase Light phase

Phase path2)

Heavy phase

Light phase

S Separator M Mixer ~

~-

Single-stage extractor data: Drum volume: 0.003-0.12 m3 Revolutions: 6500-4400 min-' Flow capacity: 1.25-120 m3/h Data of a BXP-Robatel separator: Volume: 0.017-0.22 m 3 Revolutions: 2900- 1000 min-' Drum diameter: 0.32-0.8 m Flow capacity: 6-50 m3/h

6.2 Liquid-Liquid Extraction

455

Table 6-11. (Continued) Extractor type Design

Principles Characteristic data

Robatel extractor, BXP3) [6.71]

Chamber separator which may be connected to a countercurrent flow unit. Phase mixing by centrifugal mixers at the bottom of the rotating drum, phase separation in the above separation chamber, overflow of phases via a weir system with channels to reach the next stage. Data of a BXP-Robatel separator: Volume: 0.017-0.22 m3 Revolutions: 2900- 1000 min-' Drum diameter: 0.32-0.8 m Flow capacity: 6-50 m3/h

Podbielniak centrifugal extractor') [6.76]

A series-connection of concentric perforated cylinders. The heavy phase is charged via a central shaft and passes, driven by centrifugal forces during rotation, to the outside while the light phase is led inward. Mass transfer occurs during intensive contact of the phases in the cylinder holes. Both phases flow countercurrently. The spacing between the cylinders acts as a phase separation zone. The number of cylinders connected in series corresponds to the number of mixing and separation stages.

Data of Podbielniak extractors: Rotor volume: 0.6-983 L Revolutions: 10000-1600 min-' Flow capacity: 0.23- 132 m3/h Separation stages: 3-5 per apparatus Density difference: A@2 50 kg/m3

')

*)

3,

Representation according to MULLER[0.1, Vol. 21. Representation according to BRUNNER[6.20]. Representation according to GEBAUER, STEINER, HARTLAND [6.71].

456

6 Extraction

imentally under condition as close as possible to the operating conditions. Table 6-11 illustrates the assembly and operating modes of selected centrifugal extractors and gives some geometrical and operational data. Centrifugal extractors are noted by their small space requirement, small holdup and hence short residence times, fast phase separation, and high throughput. These apparatus are particularly suitable for the treatment of substances which have a tendency to form emulsions, with an unstable key component, and also for small phase density differences. The main disadvantages are the high investment costs, maintenance costs, and energy expenditure.

6.2.5 Selection and Design of Extraction Apparatus Table 6-12 shows a strategy diagram for the solution of an extraction problem. If the substance system, type of operation, process variations, and required number of separation stages are fixed, the type of extractor can be selected. Figures 6-45 to 6-47 give an indication of the selection. However, for the final selection and design, experiments under conditions as close as possible to the operating conditions, and scaleup knowledge are required. Figure 6-48 gives a brief summary of the separation efficiency of columns, with in-

Table 6-12. Strategy to solve an extraction problem * 0 0

0 0 0

0 0

0

0

0 0

0

Analysis of extraction problem (& x,, xu) Selection of solvent (Chapter 6.2.2) Determination of operating mode, operating conditions and process Analysis of distribution equilibrium (Chapter 1.4.2) and further physical properties (e, q, CJ etc.) Balancing, determination of total throughput, determination of ya from optimization considerations of extraction requirements and solvent regeneration, determination of solvent ratio (Fig. 6-16) Choice of disperse and continuous phase (Chapter 6.2.3.4) Calculation of the required number of theoretical separation stages by graphical or arithmetical means, determination of the theoretical extractor height (length) by the aid of the HTU/NTU concept (Chapter 6.2.3) Selection of extractor type (Chapter 6.2.5) Experimental investigation of fluid mechanics and mass transfer under conditions as close as possible to operating conditions (droplet formation and size, droplet size distribution, droplet movement, coalescence, separation behavior, mass transfer, HETS or HTU with respect to surface (interfacial) tension and mixing effects), selection of the optimum operating point (Chapter 6.2.4) Transfer of experimental results to the production unit by considering experimentally gained scale-up experience Extractor geometry (cross-sectional area, height or length, opening ratio, size of internals) and energy input (Chapter 6.2.3 or 6.2.4) Determination of the optimum operating point (mixer or rotor revolutions, pulsation intensity etc.) (Chapter 6.2.3 or 6.2.4)

* The presented order has not to be followed strictly. Sometimes iterative loops are required.

6.2 Liquid-Liquid Extraction

457

L-J Process

Separator, centrifugal extractor

holdup

b' form emulsion, poor separation

limited

Lowno.of theoretical stages required

Separator, Mixer-Settler cascade with separation aid Centrifugal extractor, Graesser, RDC, ARD

yes

1

r

Mixer-Settler cascade Centrifugal " extractor Small floor space

All column types Centrifugal extractors

theoretical stages required Graesser

I

no Small floor space

I

1

yes r

-'

Pulsed sieve-

High throughput

extractor

I I

I

no

throughput

Fig. 6-45. Selection of extractor types. Representation according to BRANDT, REISSINGER, S C H R ~ E[6.31]. R

ternals reducing the cross-sectional area as a function of the total cross-sectional load. Figure 6-49 gives a comparison of the capability of selected extraction apparatus with standard dimensions, for the system toluene/acetone/water.

C=

I

I

load range Wide load range

Pulsed packed column Karr column Scheibel (OldshueRushton)

The optimal apparatus is the apparatus which best solves the extraction problem with minimum investment costs. In Fig. 6-49 the course of the characteristic cost value for different extractor types C, is presented

costs ($) throughput (m3/h) number of separation stages

(6-130)

458

6 Extraction

Fig. 6-46. Qualification of different extractors according to data of the companies Podbielniak and Luwa. Representation according to MULLER[0.1, Vol. 21.

This is equivalent to the investment costs according to STICHLMAIR [6.77]for different types of extraction apparatus, over the total throughput. The design of an extractor usually begins with experiments in a pilot-scale plant. The present correlations for the phase flow, coalescence behavior, mass transfer, etc., for different extractor types, are functions of the apparatus geometry, operating conditions, and substance properties. They mainly describe the behavior in observed experiments. Coalescence times, number of stages and height, flooding point, holdup, and optimal operating conditions (optimum value of the residence time) are found experimentally. They are used with applicable scale-up methods from the manufacturing companies, to design and dimension the full-scale extractor.

6.3 Solid-Liquid Extraction (Leaching) Solid-liquid extraction [0.1, 0.8, 6.20, 6.77-6.801 is mainly applied as percolation extraction in cross-flow and countercurrent flow and as immersion extraction in discontinuous or continuous modes. With percolation extraction, the mechanically ground and decomposed solid is moved through the extraction apparatus from stage to stage and sprayed with solvent. In this case, the solvent is enriched with the key component. Percolation extraction may only be applied to solids which allow the throughflow of solvent in a packed bed state. When the permeability of the solid is too low, immersion extraction is used. Here the solid is suspended in the solvent or is intermediately

6.3 Solid-Liquid Extraction (Leaching)

459

extracted and then separated from the enriched liquid phase, either discontinuously or by using a decanter, or continuously using special extraction apparatus. Solid-liquid extraction can be distinguished from liquid-liquid extraction by the following characteristics : 0

Fig. 6-47. Qualification of different extractors. The size of rectangle in each field is a scale of how good the extraction apparatus fulfills the quality factor. Representation according to MULLER[0.1, Vol. 21. Quality factors 1 Investment costs 2 Operating costs 3 Possibility of high stage number 4 Possibility of large throughput 5 Flexibility (reliable operation at fluctuating operating conditions) 6 Low holdup of both phases 7 Low holdup of either one of the two phases 8 Required floor space 9 Required structural height 10 Ability to treat system with emulsion formation tendency 11 Difficulties to scale-up from experimental scale to production scale

Extrmtor types I I1

I11 IV

V VI VII VIII IX

X

Spray column Packed column Sieve tray column Rotating disc contactor Mixed column Graesser contactor Pulsed column Mixer-Settler cascade Mixer-Settler tower Centrifugal extractor

0

No defined distribution coefficients exist for the distribution of the valuable or key component in the feed and solvent phase. A true equilibrium state is hardly ever reached since the solid always contains undissolved key component inside the capillaries and the enriched solvent contacts the solid surface with different key component loadings. An apparent equilibrium is reached when the solution inside the capillaries is of the same concentration as the free solution. The time required to reach this apparent equilibrium is a decisive function of the type, particle size and porosity of the solid, the solubility of the solvent, and the temperature. Usually, a temperature increase favors solid-liquid extraction. Some of the solvent, the “bonded solution”, remains bonded to the solid particle surface due to adsorption. The higher the percolation velocity, the lower the fraction of bonded solution.

Solid-liquid extraction may also be presented in a simplified equilateral triangle. With cross-current leaching, fresh solvent is added to the solid stagewise. For example, in the first extraction stage, fresh solvent is added to solid with an initial concentration w, at the state point F (Fig. 6-50). The mixing point P, lies at the intersection of the line FL with the line SP, representing the chosen solvent ratio v = i/f In the first extraction stage, the extract phase has the state point Q1. (Q, lies at the intersection of the extension of TP, and the triangle side KL,according to the separation of the mixture into the pure carrier P, and

460

6 Extraction

a1

I

bl

n, [rn-

0

10

20

30

6 [rn3/(m2-hl]-

50

LO

Fig. 6-48. Maximum efficiency of different extractor types as a function of total load. Representation according to PILHOFER, S C H R ~ E[6.46]. R

a) System: Toluene/Acetone/Water Symbol Extractor type ps[ 0701 U

V I 0

+

0

MCE Kuhni column

PSE PSE

PSE PSE

10.8 11.8

22.0 33.0 39.0

50.0

b) System: Toluene/Acetone/Water n EC column A SHE column


0.67 0.67 0.57 0.57 0.57 0.57

c) System: Butyl acetate/Acetone/Water < / p d = 0.89. Symbol Extractor type p,[o/o]

+

0

0

*

ps Z,

n,

10.8 11.8 22.0 33.0 39.0 50.0 50.0 50.0

PSE PSE PSE PSE PSE Karr column Opening ratio Cell height, tray spacing Theoretical stages per m

I 0

MCE Kiihni column

2, [mm]

75 70

100 100 100 100 50 50

6.3 Solid-Liquid Extraction (Leaching) 20

a)

t 6

10

n,[rn-’]

L

2

0.6

0.L 0.2 1

Fig. 6-49. Efficiency and characteristic cost value C of different extractor types. Representation according to STICHLMAIR [6.70]. a) Efficiency data of investigated extractor types d + - c System : Toluene/Acetone/Water Proportion by volume: = 1.5 b) Cost value C as a function of the total throughput for QVF standard extractors d e c System: Toluene/Acetone/Water = 1.5 Proportion by volume:

F$/c

1

2

L 6 10 20 6 [m3/(m2.h)]

-

L O 6 0 100

Total throughput [rn3/h]

vd/c

-

Characteristic data of investigated extractors: Type

Diameter (mm)

Graesser PSE CE Karr PFK PC SE MS Kuhni RDC

100 50 150 50 70 70 75 100 150 70

Active length (m)

1 .o

3 .O 2.0 2.6 3.0 3.0 2.0 1.2 2.0 1.7

Free cross section

Hole size

Cell length

(070)

(mm)

(cm)

-

-

20 10.8 53

2

2.5 10 7.5 2.5

-

3

-

11.8

-

461

-

9

-

2.3 -

-

15; 25; 45; 60

-

462

6 Extraction

maining wet product M, and the extract phase Q,. The remaining wet solid of the second stage is now treated with fresh solvent in the third stage, and so on. Finally the remaining wet solid is obtained with the state point M, which has a residual concentration w, of the key component according to R, after separation of the moisture. To determine the required number of extraction stages the construction follows the above procedure described until the last line M,L is below the line Therefore, the number of required extraction stages Nt is the number of points Q,, QZ, . .., Q, on the triangle side KL.In the example shown in Fig. 6-50, Nf= 2, i.e., two extraction stages are required to achieve a final concentration of key component lower than w,, in the solid. In countercurrent leaching solid is guided countercurrently to the fresh, charged solvent. During the process, the solvent becomes enriched with key component according to the process in Fig. 6-15, whose presentation is analoguous to that for liquid-liquid extraction. With a known concentration of the key component in the feed in the leaving extract phase E, and in the residual solid R , the state points F, E, and R, are given. Hence, the extraction pole P lies at the intersection of the exten-

m.

Fig. 6-50. Graphical determination of required number of stages in cross-current leaching.

the liquid phase Q , which consists of the key component S and solvent L). After separation of the solid and solution, only a certain amount of “bonded solution” resides in the solid. The state point M, of the wet solid in the first extraction stage then lies at the intersection of the line TQ, and the curve OK, found experimentally and representing the amount of bonded solution (residual solution in the solid). In the second extraction stage, fresh solvent is added to the wet solid of M , . The resulting mixture represented by point P, consists of the reK

Fig. 6-51. Graphical determination of number of stages in countercurrent leaching.

6.4 High Pressure Extraction (Distraction)

sions of the lines E,F and R,L (Fig. 6-51). Assuming a quasi equilibrium in the form of a concentration equalization between the free and bonded solution, for the process in Fig. 6-51, the required number of separation stages Nt may be determined. Nt is then the number of extract phase state points E,, E,, E,, . . ., En on the triangle side KL. (The description of the procedure [6.77] to obtain the extraction pole is found in Chapter 6.2.3.4 and for the procedure to find the required number of separation stages, see cross-current leaching). Balancing over the solid-liquid extraction stages to determine the solvent requirement, extraction yield, and number of stages is analogous to that for liquid-liquid extraction. Further information for the design of solid-liquid extractors is given in [0.8, 6.771. The final design of the extractor and the operating conditions can only be found experimentally and with appropriate scale-up knowledge. Some design forms of solid-liquid extractors are presented and characterized briefly in Table 6-13. Figure 6-52 shows the flow diagram for an oil seed extraction plant, to produce sweet oil, lecithin, and high protein meal. ~

~

6.4 High Pressure Extraction (Distraction) With high pressure extraction, (HPE) [6.82-6.861 highly compressed gases such as carbon dioxide, nitrogen, argon, ethylene, propane, etc., (Table 6-14) are mainly used as solvents in a supercritical state. Highboiling substances are also dissolved:

Removal of pollutants or unwanted components from the product by extraction (for example, to remove caffeine from coffee beans, separation of nicotine from tobacco)

a

a

463

Extract production, if valuable substances are separated from a raw product and the feed carrier is a valueless raffinate phase (for example, hop and spices extraction, oil seed extraction, valuable substance extraction from coal or from residues of carbon processing) Separation of raw material in usable residues and extracts (for example, flavor/fat separation in cocoa butter, oil/lecithin separation from soy lecithin) [6.87]

HPE is operated under pressure from 50-300 bar and in the temperature range 10- 100 "C above the critical temperature of the solvent. Gentle thermal separation of substances with high boiling points is achieved. In the application of HPE for the processing of food and luxury goods, carbon dioxide is used as the important solvent. Carbon dioxide exhibits good loadability for key components, is physiologically safe, and allows gentle thermal separation under easily controllable pressures due to the favorable physical properties (see Table 6-14a). Figure 6-53a shows solubility curves for caffeine in carbon dioxide. The solubility of supercritical gases is a function of pressure and temperature and is influenced by auxiliary components (HPE phase equilibria [6.88-6.901). Hence, by means of suitable organic components, the solubility may often be increased [6.88]. However, other mainly inorganic components such as nitrogen (Fig. 6-53 b) reduce the loadability and are only of interest with extract separation and to increase the selectivity of separation processes [6.91]. In the simplest case, an HPE unit consists of an extractor and a separator to separate the extract phase. In the extractor the supercritical solvent mainly dissolves the product (key component) under high pressure and low temperature and to release it in the separator, under low pressure and/or

464

6 Extraction

Table 6-13. Solid-liquid extractors. Extractor type

Principles, Characteristics

Cascade extraction unit' a) Charging, b) Discharging,

c) Phase separation

Solid f e e d .

+Extract

7

r, I

Cascade extraction or enrichment process, countercurrent flow of solvent in individual vessels, percolation extraction, discontinuous operation with respect to solids mode.

I I I

I I J

Drive

Extraction resldue (spent solid1

Q+- -+-Solvent

Bollmann basket conveyor extractor' Hildebrandt screw conveyor extractor' Solid feed Solid f e e d

c

Extracted

0 0 13 0

Continuous percolation extractor' Solvent I

+Extractjve.

Feed solid I

1 r - 7 --. I -i t

Exit

Spent rollds with remaining solvent Key component

--

Key component concentration

cSolids lerlractive1 path

1

In a Bollmann extractor solids are continuously passed through, percolation extractor, gas-proof shell with endless chain and buckets for solids transport, cocurrent flow of solids and sprayed solvent on the left side, countercurrent flow on the right side of the bucket elevator, self-filtration effect. The Hildebrandt extractor operates according to the dipping principle, immersion extractor, body contains two vertical and one horizontal screw, countercurrent flow of solvent and solids. While the solids pass through the unit they are compacted by the screws; extract phase is discharged via rod cages.

Percolation extractor with solids transport by a belt conveyor from chamber to chamber countercurrently to the solvent. Solvent is sprayed o n the solids by means of feed pumps, increase of extraction yield possible by increasing solvent supply and if the solvent is circulated before it is passed to the next stage.

6.4 High Pressure Extraction (Distraction)

465

Table 6-13. (continued) Extractor

type

Principles, Characteristics Revolving extractor, countercurrent gravity percolation extractor, solventproof body with slowly rotating vanes (rotating individual extraction sections). The rotary-vane passes the continuously fed conditioned solids from the feed shaft over a fixed sieve tray in one rotation to the solids discharge shaft. Sieve tray slots are concentric and expanded at the bottom. At the same time solvent is trickled over the solids. Solvent percolates through the solids and the sieve tray and collected in underneath extract chambers. Countercurrent flow of solids and solvent. Depending on the efficiency revolving extractors may be operated in single or multiple units mounted on top of each other, single or multiple solids passes. Throughput: 2-2000 tons/d solids.

Revolving extractor*

i t

-

I Spent

Extract phase Key component content

'Solids

4

I

Solvent inlet

exit

Solid path

Solvent inlet

Solid

content in solvent

Liquid path

--

I

Extract phase outlet

I

+jUSJ'-4 I

--I 1 i

Belt extractor, Lurgi'

I':

i s d i d feed

r+

~

w

~-

--

Belt extractor with endless belt supported by sieve belts for solids conveying, solids are fed to the upper belt, after passing through the upper extraction chamber, solids are given to the lower belt and discharged via a draining zone, solvent flows countercurrently to solids, solvent discharge via a filter stage*

b - ----'

Solvent inlet -

Representation according to VOESTE,WESP [0.1, Vol. 21. Representation according to Extraktionstechnik, Hamburg, Germany.

466

6 Extraction

r

Direct extraction

Group 000 Seed grinding or grooving

Group 100 Rough-pressing (Expeller) ~

:

Expressed oil Refining

6.4 High Pressure Extraction (Distraction)

467

Table 6-14. Critical data (a) and magnitude of physical characteristics (b) of some gases used in high pressure extraction units [6.85].

(a) Critical data Extraction solvent

Carbon dioxide Nitrogen Ethylene Propane Toluene

Critical data

31.3 .147.2

73.8 33.9 50.8 40.6 41.6

9.5

96.8 320.8 ~

~

0.45 0.31 0.28 0.22 0.26

~

(b) Magnitude of physical characteristics of gases Property

State Gaseous

Density (kg/m3) Diffusion coefficient (m2/s) Viscosity (Pa . s)

1

10-5

Liquid

Supercritical

1000 5 10-10

300

f

lo-' 10-5

4 Fig. 6-52. Flow chart of an oil seed extraction plant to produce sweet-oil, lecithin and meal.

Representation according to Lurgi GmbH, FrankfuWMain, [6.81]. TDC Toaster-dryer-cooler. ALCON ALCON process for preparation and/or rough-pressing

468

30

6 Extraction

LO

50

-

60

9 ["C]

70

80

Fig. 6-53. Solubility of caffeine in CO, as a function of temperature (a) and in C 0 2 / N 2 mixtures as a function of pressure (b). Representation according to GAHRS[6.97].

raised temperature. The key component may also be transferred to an auxiliary substance which ad- or absorbs the key component. The regenerated or product free solvent is now recycled, since the desired final concentration of key component is reached in the feed. By adding a suitable substance before entering the separator, for example nitrogen or argon, the solubility of the solvent is reduced, and the HPE may be operated with corresponding cost advantages resulting from these isobaric and isothermal operation conditions. Figure 6-54 shows a simplified flow sheet for a single stage, high pressure extraction plant for aromatics, including an explanatory block diagram.

Table 6-15 shows some applications of high pressure extraction under typical process parameter ranges. HPE may also be applied in multistage processes with countercurrent phase flow. Different products are obtained by a stagewise pressure release of the enriched solvent in the gaseous state (separation of oleic acid and stearic acid in a multistage process using supercritical ethylene as a solvent [6.93], sweet oil extraction and raffination [6.94], removal of caffeine from raw coffee [6.95], treatment of waste oil [6.96], etc.).

6.4 High Pressure Extraction (Distraction)

Raw material \, to I extraction ’ process ,’

Conditioning and charging

Filling i f unit

Working tank

j

__

7

Circulation and heating to supercritical extraction conditions

’

discharge

Fig. 6-54. Flow sheet of a high pressure extraction plant for aromatics including block diagram. Representation according to EGGERS and TSCHIERSCH [6.86]. A Extractor B Separator F Filter S Sorption vessel R COz recovery and working tank PI Pressure buildup pump p2 Circulation pump p3 Pressure pump H Heater Condenser

<> <3 GO, Recovery

Spent material

469

470

6 Extraction

Table 6-15. Application of high pressure extraction (a), operating parameter and product yield (b). Representation according to GAHRS[6.92].

Task

Example

Product

Remark

Product refinment

Coffee Tea Tobacco Raw lecithin (soybean oil fraction) Collagen

Caffeine-free Nicotine reduced

Production unit selective by N,/CO, mixture continuous processing possible

Aroma/flavor recovery

Hop Spices Roses

Stable extract Production unit Partialy fractionated Production unit product form

Color recovery

Paprika Orange peel

Flavor-free natural dye

Active substance recovery

Valerian Camomile Drugs

Active substance concentrate

Oil recovery/ refinment

Jojoba, Rape and other substances

High value extract (color, scent, flavor)

Free-flowing Fat-free

(b)

Substance

Extraction conditions P L9 (bar) (“C)

Gas load

Yield

(g/kg)

(070)

HOP Nutmeg flower

200-400

40-50

20-30

10-15

250-350

40-60

2-3

10-12

Paprika *

200

Pepper *

.........

14-15

40-60

300-400

3-5

80- 120

3-5

.........

40-50

.........

35-40

25-35

4-6

200-250

Clove

*

Fractional extraction.

.........

100- 200

15-18

......... 2-3 6-8

......... 8- 10 15-21

References

References [6. I] GROB,K. : Fliissig-Extraktion. DECHEMA-Erfahrungsaustausch, Frankfurt/ Main 1954. [6.2] ALDERS,L. : Liquid-Liquid Extraction. Elsevier Publishing Cornp., Amsterdam 1959. [6.3] T ~ E Y B A L R., E.: Liquid Extraction. McGraw-Hill Book Cornp., New York 1963. [6.4] SHERWOOD, T. K., and PIGFORD,R. L.: Absorption and Extraction. McGrawHill Book Comp., New York 1952. [6.5] HANSON, C. : Neuere Fortschritte der Fliissig-Fliissig-Extraktion.Sauerlander, Aarau 1974. [6.6] HARTLAND, S. : Countercurrent Extraction. Pergarnon Press, Oxford 1970. [6.7] MULLER, E. : ,,Flussig-Flussig-Extraktion". Ullmanns Encyklopiidieder technischen Chemie, Vol. 2, 4th Ed. Verlag Chernie, Weinheirn 1972. [6.8] ~ E Y B A L , R. E.: Liquid Extraction in Chemical Engineers Handbook. McGrawHill, New York 1973. [6.9] Lo, T. C., BAIRD,M. H. I. and HANSON, C. : Handbook of Solvent Extraction. J. Wiley & Sons, New York 1983. [6.10] SOUDERS,M.: Chem. Eng. Prog. 60 119hAl ,-- 7 75-. [6.11] HAMPE,M.: Chem. Ing. Tech. 50 (1978) 9, 647-655. K. : Chem. Tech. (Heidelberg) [6.12] LACKNER, 11 (1982) 5 , 497-502. [6.13] HAMPE,M.: Chem. Zng. Tech. 57 (1985) 8, 669-681. T.: Chem. Ind. 35 (1983) 3, [6.14] PILHOFER, 156-157. [6.15] NITSCH,W. : Fortschr. Vefahrenstech. sec A 20 (1982) 59-70. [6.16] RITCEY,G. M.: Sep. Sci. Technol. 18 (1983) 14/15, 1617-1646. [6.17] HUNTER,T. G., and NASH,A. W.: J SOC. Chem. Ind. 53 (1934) 95 T. K., BAUERMANN, H. D., [6.18] BAUCKHAGE, BLASS,E., SAUER,H., STOLTING,M., J., and WAGNER,H.: T~NHUMBERG, Chem. Ing. Tech. 47 (1975) 5, 169-222. ' I

-1

'

471

[6.19] BLASS, E., GOLDMANN,G., HIRSCHMA", K., MIHAIIDWITSCH, P., and PIETZSCH,W.: Chem. Zng. Tech. 57 (1985) 7, 565-581. [6.20] BRUNNER, K.-H. : ,,Separatoren und Dekanter fur die kontinuierliche Extraktion," Tech. Wiss. Dok. 4 (1982). Westfalia Separator AG, Oelde. [6.21] BAUER,R.: Dissertation, ETH Zurich 1976. [6.22] BORNTON, J. D.: Trans. Znst. Chem. Eng. 35 (1957) 316-330. J., and [6.23] REISSINGER, K.-H., SCHR~TER, BACKER,W.: Chem. Ing. Tech. 53 (1981) 8, 607-614. [6.24] MARR,R. : Chem. Zng. Tech. 50 (1978) 5, 337-344. [6.25] STEMERDING, S., and ZUIDERWEG, F. J. : Chem. Eng. London 5 (1963) 156. [6.26] MARR,R., and MOSER,F.: Chem. Ing. Tech. 50 (1978) 2, 90-100. [6.27] GAYLER,R., ROBERTS,N. W., and PRATT,H. R. C.: Trans. Inst. Chem. Eng. 31 (1953) 57-58. [6.28] BORNTON, J. D.: Trans. Znst. Chem. Eng. 29 (1951) 89. [6.29] MERSMANN, A.: Chem. Ing. Tech. 52 (1980) 12, 933-942. [6.30] STROBEL,W., and SALZER,G.: Wiss. Z./Tech. Hochsch. Chem. Carl Schorlemmer Leuna-Merseburg 10 (1968) 39. H. W., REISSINGER, K.-H., and [6.31] BRANDT, SCHR~TER, J.: Chem. Ing. Tech. 50 (1978) 5 , 345-354. [6.32] GLASSER, D., ARNOLD,D. R., BRYSON, A. W., and VIELER,A. M. S. : Miner. Sci. Eng. 8 (1976) 23-45. [6.33] BOHNET, M. : Chem. Ing. Tech. 48 (1976) 3, 177-264. T. : Chem. Zng. Tech. 46 (1974) [6.34] PILHOFER, 18, 783. T.: Chem. Ing. Tech. 48 (1976) [6.35] PILHOFER, 3, 237. [6.36] BILLET,R., and MACKOWIAK, J.: Verfahrenstechnik 15 (1981) 12, 898-904. [6.37] BILLET,R., HUSUNG,G., and MACKOWIAK,J. : Chem. Tech. (Heidelberg) 8 (1979) 1, 25-29. [6.38] BILLET,R., and MACKOWIAK, J.: Chem. Ing. Tech. 57 (1985) 1, 56-57.

472

6 Extraction

[6.39] PRATT,H. R. C.: Znd. Chem. 31 (1955) 505 -510. [6.40] MERSMANN, A.: Chem. Zng. Tech. 52 (1980) 12, 933-942. [6.41] STREIFF,F. A.: Chem. Ing. Tech. 55 (1983) 5, 398-399. [6.42] PILHOFER, T. and MEWES,D. : Siebbodenextraktionskolonnen. Vorausberechnung unpulsierter Kolonnen. Verlag Chemie, Weinheim 1979. W., [6.43] BENDER,E., BERGER,R., LEUCKEL, and WOLF, D.: Chem. Ing. Tech. 51 (1979) 3, 192-199. [6.44] BERGER,R., LEUCKEL,W., and WOLF, D.: Chem. Ing. Tech. MS 602/78, Synopse Chem. Ing. Tech. 50 (1978) 7. H. W., REISSINGER, K.-H., and [6.45] BRANDT, SCHROTER, J. : Vkflahrenstechnik8(1975) 8 , 5. [6.46] PILHOFER, T., and SCHR~DER, J.: Chem. Zng. Tech. 56 (1984) 12, 883-890. [6.47] SCHAFER,P. : Graduierungsarbeit, FH Koln 1977. [6.48] BERGER,R., and WALTER,K.: ,,Belastungsgrenzen pulsierter Siebbodenextraktoren". Manuscript submitted to Chem. Eng. Sci. (1985). [6.49] REISSINGER,K.-H., and MARR, R.: Chem. Ing. Tech. 56 (1984) 7, 558-559. [6.50] PIETZSCH,W., and BLASS,E.: Chem. Zng. Tech. 57 (1985) 10, 872-873. E., and VOGELPOHL, A.: [6.51] AUFDERHEIDE, Chem. Zng. Tech. 56 (1984) 9, 724-725. [6.52] MARR,R.: Chem. Zng. Tech. 50 (1978) 5, 337 - 344. 0.: 16.531 MARR,R., MOSER,F., and HUSUNG, Chem. Ing. Tech. 46 (1974) 5, 207. [6.54] MARR,R., HUSUNG,G., and MOSER,F. : Chem. Zng. Tech. 47 (1975) 5 , 203. G., and MOSER,F. : [6.55] MARR,R., HUSUNG, Chem. Zng. Tech. 48 (1976) 3, 245. [6.56] WOLSCHNER, B., SOMMEREGGER, E., and MARR,R. : Chem. Ing. Tech. 52 (1980) 3, 277. [6.57] MISEK,T., and MAREK,J.: Br. Chem. Eng. 5 (1970) 2, 202-207. [6.58] ZEHNDER,W. H.: In$ Chim. 8 (1970), 3/4. [6.59] MARR, R.: Chem. Tech. Heidelberg 5 (1976) 9, 373-376.

[6.60] SCHEIBEL,E. G.: Ind. Eng. Chem. 42 (1950) 1048-1057. [6.61] BONNET,J. C., and JEFFREYS, G. V.: AIChE J. 31 (1985) 5, 788-801. [6.62] OLDSHUE,J. Y,, and RUSHTON,J. H.: Chem. Eng. Prog. 48 (1952) 297. [6.63] MOGLI,A.: Chem. Ing. Tech. 37 (1965) 3, 210-213. T. : Auslegungskriterien von [6.64] PILHOFER, Ruhrzellenextraktoren wie dem Graesser-, dem Kiihni- und dem QVF-Extraktor. Company report from Fa. QVF Glastechnik GmbH, Wiesbaden. [6.65] STEINER,L., and HARTLAND, S.: Chem. Rundsch. 33 (1980) 13, 1-5. [6.66] Company report from: QVF Glastechnik GmbH, Wiesbaden. [6.67] GAUBINGER,W., HUSUNG, G., and MARR, R. : Maschinenmarkt 89 (1983) 35, 758-760. [6.68] SHEIKH,A. R., INGHAM,J., and HANSON, C.: Trans. Inst. Chem. Eng. 50 (1972). L6.691 HUSUNG, G. : Chem. Zng. Tech. 56 (1984) 7, 548-549. [6.70] STICHLMAIR, J.: Chem. Zng. Tech. 52 (1980) 3, 253-255. [6.71] GEBAUER, K . , STEINER,L., and HARTLAND, S.: Chem. zng. Tech. 54 (1982) 5, 476-496. [6.72] S ~ L T I N GM., , HENGSTLER,N., and BLASS,E.: Ber. Bunsenges. Phys. Chem. 83 (1979) 1116-1120. [6.73] S ~ L T I N G M., : Dissertation, TU Miinchen 1979. [6.74] SCHILP,R., and BLASS,E. : Chem. Eng. Commun. 28 (1984) 1/3, 85-98. F., and BLASS,E.: Chem. [6.75] OTILLINGER, Zng. Tech. 57 (1985) 9, 796-797. [6.76] TODD,T. B. : Chem. Eng. Prog. 62 (1966) 8, 119-124. K., and VOESTE,T.: [6.77] SCHOENEMANN, Fette, Seifen, Anstrichm. 54 (1952) 358. H. P. : Neuzeitliche Techno[6.78] KAUFMANN, logic der Fette und Fettprodukte. Aschendorffsche Verlagsbuchhandlung, Miinster 1959. [6.79] SWERN,D.: Bailey's Industrial Oil and Fat Products. 15/4. Solvent extraction. Interscience Publishers, New York 1964.

References 16.801 EIGLM~LLER, J. K.: Osterr. Chem. Ztg. 56 (1955) 56, 221. [6.81] Company report: ,,Kontinuierliche Liisungsmittel-Extraktion von Olsaaten“, T 1426 5 (1982), Fa. Lurgi, Frankfurt. [6.82] ZOSEL,K.: Angew. Chem. 90 (1978) 748-755. [6.83] SCHNEIDER,G. M., STAHL,E., and WILKE,G. : Extraction with Supercritical Gases. Verlag Chemie, Weinheim 1980. [6.84] BRUNNER, G., PETER,S.: Chem. Ing. Tech. 53 (1981) 529-542. [6.85] GAHRS,H. J.: Znt. Z. Lebensm. Technol. Verfahrenstech. 35 (1984) 4. [6.86] EGGERS, R., and BCHIERSCH, R.: Chem. Zng. Tech. 50 (1978) 11, 842-849. [6.87] COENEN, H., EGGERS,R., and KRIEGEL, E.: Chem. Tech. (Heidelberg) 11 (1982) 10, 1199-1204. [6.88] BRUNNER, G.: Fluid Phase Equilib. 10 (1983) 289. 16.891 SCHNEIDER, G. M . : Angew. Chem. 90 (1978) 762-774. [6.90] STEPHAN, K., and SCHABER, K.: Chem. Zng. Tech. 50 (1978) 9, 718-719. [6.91] GAHRS,H. J.: Ber. Bunsenges. Phys. Chem. 88 (1984) 894. [6.92] GAHRS,H. J.: Chem. Tech. (Heidelberg) 14 (1985) 7, 51-55. G., and RIHA,R.: [6.93] PETER,S . , BRUNNER, Dechema Monogr. 73 (1974) 197-204. [6.94] COENEN, H., and KRIEGEL, E., (Krupp), D. B. P. (0s) 2843920 (1978). [6.95] COENEN, H., and KRIEGEL, E., (Krupp), D. B. P. 2846976 (1978/1981).

473

[6.96] COENEN, H., and KRIEGEL, E., (Krupp), D. B. P. 2850540 (1979/1982). [6.97] GAHRS,H. J.: Chem. Znd. 37 (1985) 7, 477-479. 16.981 SCHREINER, H.: Chem. Ztg., Chem. Appar. 91 (1967) 18, 667-676 and 93 (1969) 24, 971-982. (6.991 Lo, %H, C., BAIRD,M. H. I., and HANSON,C.: Handbook of Solvent Extraction. Wiley, New York 1983. [6.100] ,,Essigsaure-Ruckgewinnungmit Flussig-Flussig-Extraktion“, Chem. Tech. (Heidelberg) 11 (1982) 12, 1392- 1393. [6.101] BAILES,P. J., GLEDHILL, J., GODFREY, J. C., and SLATER, M. J. : Chem. Zng. Tech. 58 (1986) 10, 807-809. [6.102] REISSINGER,K.-H., and MARR, R.: Chem. Zng. Tech. 58 (1986) 7, 540-547. [6.103] BILLET,R., BRAUN,C., MACKOWIAK, J., and PAJAK,M.: Chem. Zng. Tech. 58 (1986) 6, 488-490. [6.104] BLUMBERG, R. : Liquid-Liquid Extraction. Academic Press, London, San Diego 1988. E . , BERGER, R., KOSTERS, C. G., [6.105] MULLER, and Cox, M. : “Liquid-Liquid Extraction”. Ullmann’s Encyclopedia of Zndustrial Chemistry. VCH Verlagsgesellschaft, Weinheim 1988. [6.106] COULSON,J. M., RICHARDSON, J. F., BACKHURST, J. R., and HARKER,J. H.: “Liquid-Liquid Extraction”. Chemical Engineering. Vol. 2. Pergamon Press, Elmsford, Oxford 1991.

Thermal Separation Processes: Principles and Design Klaus Sattler, Hans Jacob Feindt copyright 0VCH

Verlagsgesellschaft m h H , 1995

7 Solvent Evaporation, Crystallization

7.1 Basic Concept and Processing Modes of Crystallization Crystallization is a thermal, phase-forming separation process: from an amorphous phase, at least one solid crystalline phase is formed by phase transformation or by crystallization from a solution, which has one or more solids molecularly dispersed in a solvent, or by crystallization from a melt, or by desublimation out of a vapor phase. A crystalline product is produced from an initial amorphous phase, liquid or gas, with one or more components, for the purpose of forming, cleaning or mixture separation. The produced crystals usually have to be cleaned and require aftertreatment as shown schematically in Fig. 7-1. The crystal size distribution and crystal form mainly depend on the operating conditions such as pressure, temperature, degree of supersaturation and concentration of separating agents influencing the yield. Crystallization from a solution is a thermal separation process. A solution of one or more solids molecularly dispersed in a solvent is usually concentrated by multistage evaporation of the solvent. Since the concentrated solution becomes supersaturated by cooling (“cooling crystallization”), evaporation of solvent (“evaporation crystallization”) or flash evaporation (“vacuum crystallization” a combination of cooling and evaporation) causes crystals to form and grow. To reduce the degree of supersaturation, the surplus forms a solid which

,

may be mechanically separated from the remaining solution. Therefore, the dissolved substance is separated from the original solution. For a certain particle size distribution formation of seeds and crystal growth are carried out in “classifying” crystalliza-

Crystallization from a solution by cooling, solvent evaporatlon or a cambination of both ---

Solid liquid separation b v mechanical means Raw crystals

Mother liauor

Filtrate (washing possibly

Pure crystals

soiled)

Solvent Product crystals

b) Crystallization from a melt (simple or fractional)

4

Crystals

a vapor phase

Crystals in a desubllmer

crystals

Product crystals

Fig. 7-1. Aftertreatment of crystals. a) Crystallization from a solution b) Melt crystallization c) Desublimation

476

7 Solvent Evaporation, Crystallization

tion apparatus, by controlling the operation parameters such as cooling rate and evaporation rate, flow control etc. Figure 7-2 shows, for example, crystallization of ammonium sulfate from a crude solution, obtained from a caprolactam process, in a three stage evaporation crystallization unit [7.16]. The unsaturated crude solution is fed to a crystal washer (l), in which

the fines in the slurry from the crystallizer ( 5 ) are dissolved, and thereby impurities are washed out. The crude solution and cleaned slurry are then charged to a slurry vessel (2) and concentrated. The thickened mixture is separated by means of a centrifuge (3) into salt and a crystal-free solution. The salt is formed after drying in a dryer (4). The solution flows to the third crystallization stage

Mother liquor outlet

F e e d inlet

t

Ammonium sulfate

Fig. 7-2. Flow sheet of a three-stage crystallization unit of ammonium sulfate from a crude solution of a caprolactam process. Representation according to Mannesmann Engineering AG, Messo-Chemietechnik. I Crystal washer 5 Messo vortex crystallizer 2 Slurry vessel 6 Evaporator 3 Centrifuge 7 Circulation pump 4 Dryer 8 Vapor condenser

Process example: Crude solution throughput Amount of crystals Water evaporated Steam consumption Temperature levels Heating temperature difference in evaporator Mean crystal particle size Specific crystal production capacity, refferred to cross-sectional area

ca. 41 t/h (37-39% ammonium sulfate) ca. 15 t/h ca. 24 t/h ca. 10 t/h (0.67 t/t product) 100 + 75 50°C ca. 6-8 "C ca. 1.5 mm +

ca. 400 kg/(m2. h)

7.1 Basic Concept and Processing Modes of Crystailization

and from there is fed backward to the second and first stages, respectively. In these stages some of the concentrated solution with additional substances is discharged. The crystallization actually occurs in classifying crystallizers ( 5 ) with external heat circulation (evaporator (6) with stagewise usage of the vapor heat, circulation pump (7)) and internal circulation to control seed formation and crystal growth. The crystal slurry obtained in the crystallizers is charged via nozzles, to the crystal washer (1) where it is again mixed with crude solution. The crystallization process from a solution, is accelerated if the solubility of the dissolved substance is decreased or made virtually insoluble, by adding a separating agent. The solute is then precipitated. With precipitation crystallization, without previous chemical reaction with the solvent, an auxiliary component (the separating agent) is added. This component has a higher solubility in the solvent than the dissolved substance, which is then forced to precipitate by the separating agent. For example, in precipitation crystallization, strong electrolytes such as salts are replaced in aqueous solutions by nonelectrolytes (Fig. 7-3 ; t7.1, 7.141). An example is the precipitation of ammonium alum, with ethanol as a precipitating agent [7.171. Another variation of crystallization out of a solution, is separation by adding a salt (chemical precipitation). Nonelectrolytes or weak electrolytes are replaced by strong electrolytes in aqueous solutions t7.1, 7.181. Adductive crystallization, crystallization from a solution by forming an adduct (inclusion complex), such as with urea and thiourea as auxiliary substances is of increasing importance especially in crude oil treatment. Urea as a solid or a solution is added to the solution to be treated. A crystalline addition compound (an adduct) is formed with a straight-chain organic compound, which may then be mechanically separated from the remaining solution. This

477

Organic

Organic precipitating

-v1 I

do

A-

Aqueous

solution

Effluent

Crystals

Fig. 7-3.Precipitation unit [7.14].

CR Crystallizer PR Precipitating agent receiver AR Precipitating agent recovery by a rectification unit, in this case the precipitating agent has a lower boiling point than water

is a simple way to separate branched and straight-chain organic compounds. Adduct formation is an exothermic process. These “loose” addition compounds are easily separated by dissociation, either thermally or by addition of solvent t7.1, 7.151. Figure 7-4 shows a process to dewax crude oil by adductive crystallization, adding aqueous urea solution and methylene chloride as activators. Adducts are formed in the mixing vessel (a); the generated spheroid adduct particles are sieved, and are treated with saturated steam and discomposed at 75°C in the mixing vessel (b), where adhered methylene chloride vaporizes and recirculates as condensate in the mixing vessel (a). The resulting phases of adduct separation, a rich wax floating on

478

7 Solvent Evaporation, Crystallization

*

C

Solvent recovery

Dewaxed oil

e

d

L-

Aqueous phase +Wax

fParaffin)

the surface, and the aqueous urea solution are separated in a separator (d). The waxrich phase is separated to wax and methylene in the column (e); the urea solution is concentrated in an evaporator (f) and recycled to the adduct-forming stage (a). The filtrate from the sieving stage is separated in the column (c) to dewax crude oil and methylene chloride, which is recycled. In spray crystallization (prilling) [7.2] hot, high temperature concentrated solutions or melts are dispersed in a tower by means of nozzles or atomizer discs into a cold air flow, wherein crystallization heat is removed by direct phase contact. If the crystallization starts after droplets are formed, smooth spherical particles, “prills” result, otherwise irregularly shaped products with a partial crystalline or amorphous structure are formed at high cooling rate and evaporation rate.

Fig. 7-4. Flow sheet of an adductive crystallization process to dewax crude oil by the aid of aqueous urea solution and methylene chloride, process of Deutsche Erdol AG. Representation according to MATZ[7.2]. a) Adductive crystallization mixing vessel b) Adduct separator c) Filtrate separation column d) Phase separator e) Column to separate wax-rich phases 9 Evaporator to concentrate urea solution g) Condenser

Freezing is a special kind of crystallization from solution. In this case, the solvent crystallizes rather than the solute, resulting in a thermally gentle concentration of the dissolved substance in the remaining solvent [7.1]. Figure 7-5 shows a simplified flow sheet of a plant which combines freeze concentration and freeze drying of products. A freeze concentrator is used as the crystallizer. The principle of freeze crystallization is shown by the solution freezing point curve. The crude solution is first concentrated in the crystallizer (1) by the freezing of ice, which is then separated from the concentrate by a centrifuge (2). The concentrate is frozen (3), grained (4), sieved (3, dried in a freeze dryer (6) and packed in weighed portions (7). Double-jacketed heat transfer tubes with rotating internal shaft act as crystalh e r s . Spring-mounted wipers on the shaft,

7.1 Basic Concept and Processing Modes of Crystallization

479

Table 7-1. Overview of crystallization processes, intermediate and final product states, control variables’). Stage

Crystallization process’) Liquid3) one-phase multicomponent system (solution or melt)

Base material

Crystallization process

Temperature increase or decrease

Control variables General control variables of the process by changing the composition of the base material: 0 Purification (physically,

Solvent removal (evaporation) 0

First intermediate product First crystallization stage Second intermediate product

Crystallization

O~ersaturated~) or subcooled one-phase multicomponent system

.iNucleation

Unstable3) (crystallizing) two (multi) phase multicomponent system

Crystal growth I

Final product(s)

Separation of final products

2,

3,

Stable’) two (multi) phase multicomponent system

Screening, filtration, centrifugation, decanting, washing

pH Level

Influence of process flow: Kinetic (process time, rate) Temperature 0 Pressure Diffusion (mixture ratio) Control of the crystal number and crystal size distribution: 0 Change of natural seed forming 0 Change of seed forming caused by process 0 Artificial seed forming (a) Additives (b) Seed crystals 0 Elimination of unwanted nucleation (incrustation) Control of crystal growth (except base materials): 0 Process programming pH Value 0 Additives

Representation according to ADAMSKI [7.19]. In comparison to mono crystal growth crystallization here refers to processes to produce various shapes and kind of crystals. Gas phase is not considered.

480

7 Solvent Evaporation, Crystallization

TI Feed

6

.....

r-

i'.

*--J i

r-'

t

( J A

L

Product

Concentrate

7

&ooltng

brine. out

Suspension

b)

n

~

Cooling brine, in

Feed

1 ) OL

0

a ["CI

4

6

-8

t

10

20

JO

40

S[%l

-

concentrate

7.1 Basic Concept and Processing Modes of Crystallization

wipe solvent crystals off of the cooling area. Multiple units are linked together, each crystallizer with a cooling area of ca. 5 m2. The refrigerant is cooling brine or a liquid cooling agent, in the temperature range of ca. -10 to -30°C. The crystallization process from a melt, containing two or more components with different melting points, is also a thermal separation process. By removing heat, the melt partially forms crystals that have a different composition to the original melt. Higher melting components are enriched in the crystals and depleted in the remaining melt. Separation of melt and crystals also implies a partial mixture separation, which may be improved by a countercurrent fractionating operation. Multistage crystallization in systems with complete miscibility of the components in the liquid and solid states and with systems which form mixed crystals, leads to an almost complete separation. In systems forming an eutectic mixture, simple separation is only possible up to the eutectic composition. A schematic for the two stage purification process of a higher melting product

from lower melting impurities, according to the Sulzer-MWB crystallization process, is shown in Fig. 7-6. A raw melt from a previous cycle with residues (R,) from tank (Tl) is collected in the tank (CT), where it is mixed with the initial melt (M,) and fed to the crystallizer (E). By partial crystallization in the crystallizer (E), M, is separated into a low-melting fraction (R,) and crystal fraction ( C , ) . (R,) leaves the crystallizer as a concentrated impurity. The cleaned product (Cl) melts in the tank (T2), together with the concentrate (C) from the previous cycle. The resulting melt (M,), acts as the initial mixture for a second crystallization stage. (M,) is separated into the low-melting residue (R,) and to the high-melting fraction (C,), from which some of it is withdrawn as pure product (C). The remainder flows into tank (T2) to be treated in the next cycle. The point is now reached for the next crystallization cycle t o begin in the same manner. The total separating efficiency depends on the number of stages, the reflux ratio and the distribution coefficient. A vertical long-tube apparatus (tube diameter 50-75 mm, tube length 12 m) is used as a crystallizer with falling film flow on the

Freeze concentration and freeze-drying plant. Representation according to Krauss-Maffei, Munich, Germany. a) Flow sheet 1 Crystallizer 2 Centrifuge 3 Freezing unit 4 Granulator 5 Screen 6 Freeze dryer I Packing 8 Condenser 9 Vacuum pump b) Crystallizer c) Crystallization principle, presented by the aid of a solution freezing-point curve d Temperature S Solids fraction

4 Fig. 7-5.

481

482

7 Solvent Evaporation, Crystallization Product cycle

Feed

Residue

Purified Droduct

bl

Feed

I R Composition

-

Fig. 7-6. Two-stage purification process according to the Sulzer-MWB crystallization process. Representation according to Sulzer MWB, Buchs. a) Simplified flow sheet E Crystallizer CT Mixing tank Pump for cooling or heating agent CP T1, T2 Stock tank CH Heat exchanger PP Product pump b) Phase diagram

side of the cooling agent and on the product side. The temperature of the cooling agent is controlled by the crystals growing on the tube wall. After crystallization, the crystals are melted and discharged. Table 7-1 gives an overview of crystallization processes from a solution and from a melt, and describes the intermediate obtained and the final products. Possibilities to control crystallization processes are also given, particularly in the stages of seed formation (nucleation) and crystal growth. These are used to influence the number of crystals, crystal size, crystal size distribution, form and structure. Crystallization from a mixed vapor phase is also a thermal separation process. By removing heat from the system, desublimable components are converted directly to crystals. The remaining gas mixture then contains lower quantities of the desublimable components; thus a partial separation is achieved by the phase separation. Desublimation may also be improved by countercurrent phase flow. Figure 7-7 shows the application of crystallization in the mixed vapor phase, to produce phthalic anhydride by oxidation of oxylene. The o-xylene/air mixture oxidizes in a tube reactor containing a packed-bed catalyst. The product mixture is then cooled in a gas cooler and raw phthalic anhydride crystallizes in a crystal separator (desublimator). The crystallizer operates automatically in “swing” mode. After crystallization, crystals are melted in the crystallizer, and then treated by distillation to obtain phthalic anhydride in the desired purity. The remaining gas from the desublimator is guided to a gas-cleaning stage. Since the melting point and triple point of substance under comparable pressure are considerably smaller than the boiling point, crystallization from a melt or from the vapor phase, provides a thermally gentle mixture separation process usually with a lower energy input than for distillation. Crystalli-

483

7.1 Basic Concept and Processing Modes of Crystallization 0)

Feed

Xvlene

Effluent gas

-cI

Heo ting

-I 1 1

I

t

Raw PA to distillation

Feed water

b)

1 3 A

Steom

Pretreatment

Heating

-

Preliminary column

E

+Final distillation

Heating

Fig. 7-7. Simplified flow sheet of a process to produce phthalic anhydride (PA). Representation according to Lurgi GmbH, Frankfurt/Main, Germany. a) Oxidation, 1'' process step b) Continuous distillation, 2"d process step

oil

484

7 Solvent Evaporation, Crystallization

zation by simple cooling of the solution is also a gentle separation process. (Special crystallization processes such as zone melting, monocrystal growth, highpressure crystallization and crystallization by transfer reaction are not discussed here. Introductory literature, for example, is given in [0.4 and 7.11).

7.2 Crystallization from a Solution The solution being treated is first concentrated, usually in a multistage concentration unit. The concentrated solution is then guided to the crystallizer where crystallization occurs. The driving force during crystallization is the concentration difference of the dissolved substance in the supersaturated solution and just saturated solution,

thus a deviation from the solution equilibrium (see Chapter 1.4.5.1). Therefore, the concentrated solution has first to be saturated and then supersaturated. Supersaturation is achieved in practice in three different ways (Fig. 7-8): With a large temperature dependency of the solubility, supersaturation is achieved by simple surface cooling of the saturated solution from 8, to tJ2,cooling crystallization If only a minor temperature dependency of the solubility is observed, the solution is supersaturated by evaporation of solvent, evaporation crystallization. The saturated solution is heated from ~9~ to 8, where the solvent evaporates along the line 2 + 3 If the solubility shows a distinct dependency on the temperature, or the solution must be thermally gently treated then solution cooling and solution evaporation are combined, vacuum crystallization. The saturated solution at state point 1 is

b)

ssc

tI

II

Fig. 7-8.Crystallization processes stages in temperature-solubilitydiagrams, a) Solution supersaturation by cooling b) Solution supersaturation by solvent evaporation c) Solution supersaturation by vacuum cooling

SC Solubility curve SSC Supersaturation curve X Dissolved solid ratio in solvent t9 Temperature

7.2 Crystallization from a Solution

485

treated under vacuum. Some of the solvent evaporates through cooling of the solution from L9, to L9,. Cooling and evaporation causes supersaturation from x,to x, With cooling crystallization, solvent supersaturation is greatest at the lowest possible temperature in the crystallizer, and, therefore, a crust forms on the cooling area. To restrict incrustation, the crystallizer must be operated with a small temperature difference between the solvent and cooling agent, which requires large cooling areas. Vacuum crystallization needs no heat transfer area in the solution, and therefore no incrustation problem occurs. To increase economic efficiency with evaporation crystallization, with respect to steam input, the process is designed for multistage operation with steam heat recovery. The first stage operates under conditions close to ambient pressure, and the final stage under a vacuum, such that the steam is condensed without compression at the temperature of the available cooling water. Once supersaturation of the solution is achieved, crystallization occurs by reducing the supersaturation until the saturation line is reached again. Small seeds are then formed which then grow to crystals (see Chapter 7.2.3). In a simplified manner Fig. 7-9 explains schematically the procedure of crystallization in solutions. The result of this crystallization is a crystalline, ready-to-sell product, whose quality is characterized by Crystal size and crystal size distribution (influencing the separability from the remaining solution, storage behavior, dust fraction, soluting behavior, free flow behavior, etc.) Crystal form (influencing the separability, residual moisture content, flow behavior, stickiness, etc.) Purity

L!/+

Crystals

Fig. 7-9. Crystallization unit, simplified flow

diagram. CR Crystallizer EU Multistage evaporation unit FU Filtration or centrifuge unit FL Feed liquor to crystallizer ML Mother liquor depending on the solid load fed back to crystallizer or evaporation stage

Sea water, residual solutions from chemical plants, salt- loaded wastewater, alkalis, alkaline-earth brines and sugar solutions are the most important raw solutions in crystallization process practice.

7.2.1 Concentration of Solutions by Evaporation Before a solution is saturated or supersaturated it usually has to be concentrated by evaporation of the solvent. Since the vapor pressure of the dissolved solid is normally negligible during evaporation, only the sol-

486

7 Solvent Evaporation, Crystallization

vent vaporizes. The remaining solid concentrates in the solution. Solvent evaporation has to be carried out with a minimum total cost. This implies minimal energy consumption with respect to the evaporation rate, but also investment costs, maintenance and service costs have to be in an acceptable range. Solution concentration by evaporation of solvent, is carried out in single or multistage evaporation units with or without thermal or mechanical vapor compression, or in multistage flash evaporation units. Common evaporation units include circulating evaporators, continuous (single pass) evaporators or flash evaporators.

D = F a . (I

-?)

DH =

d . (h,

- h,)

- Fa * (h, - h a )

h*-hw

+ Qu (7-2)

by considering

F,

= Fa - D

D.h,

I

6+ I

D H , h, Condensate

F,. h,. w, Concentrated solution

b)

Solvent vapor A

I

l i 1 - 1

Steam-

,

, 2--Dilute

solution

(7-1)

A heat balance over the evaporator gives the required steam flow rate DH to evaporate d .

r--+---

_---

7.2.1.1 Single Stage Solution Evaporation Figure 7-10 shows a schematic of a single stage solution evaporator, to evaporate a solution flow rate Fa with an initial mass fraction w, of soluted solid t o the final concentration of w,. A solvent mass balance gives the solvent flow rate d (“vapor flow rate”) to be evaporated in order to concentrate the solution from w, to w,

Solvent vapor

a)

(7-3)

I

1

Condensate

Fig. 7-10. Balance scheme to determine the solvent vapor flow and steam consumption (a) and heat flow diagram for single stage solution evaporation (b).

7.2 Crystallization from a Solution

where enthalpy of the evaporated vapor

hD

(7-4) do,

+ At9

Al9

CP.L

Ah/,, h, h,, h , hW

boiling temperature of the solution mean rise of the boiling point across the heating area (see Chapter 1.4.3.4) mean specific heat of solution evaporation enthalpy of solvent at the evaporator operating pressure p enthalpy of heating steam enthalpy of dilute and concentrated solution enthalpy of heating steam condensate

saturation temperature corresponding to heating steam pressure p H degree of subcooling of condensate

dH

Ad,

0.55 . ( 1 9~ IY,~)

(7-6)

with vertically mounted evaporator tubes and film condensation temperature on the steam side of the evaporator tube under the condensate film heat losses

19,D

Q,

The specific steam requirement dH dH=;

DH

D

(7-7)

for an aqueous solution preheated to its boiling point and fed to a single stage evap-

487

orator, is about 1 metric ton (1000 kg) of steam for the evaporation of 1 metric ton of water. Single stage evaporators are only used if a low evaporation rate with low steam costs is required, or if the generated vapor is impure, or corrosive solutions cause high equipment costs due to the use of expensive materials. For all other cases, energy saving aspects are listed in Table 7-2, and have to be considered during the design to achieve a cost-optimized solution for an individual unit suited to a particular concentration problem.

7.2.1.2 Multistage Solution Evaporation The dissolved substance is concentrated by solvent evaporation in a multistage evaporator. The aim of multistage evaporation is to reduce the specific energy costs for the evaporation of a certain amount of solvent. This is possible because the use of evaporated steam from one stage as the heating steam in the following stage, provided the operating pressure on the solvent side is reduced from stage to stage, or vapor compression is carried out. This is necessary to achieve the temperature gradient required for heat transfer between vapor and solution in each stage. The specific steam requirement d,, resulting from the ratio of the supplied amount of steam DH and total amount of solvent evaporated in the entire unit d decreases with increasing number n of connected evaporation stages. Therefore, steam operating costs C, decrease with increasing stage number n (Fig. 7-11). However, investment costs C, increase with increasing n. The optimum number of stages nopt is found at the minimum of the total cost curve C,(n)

c,= c,+ c,

(7-8)

488

7 Solvent Evaporation, Crystallization

Table 7-2. Energy saving possibilities during solvent evaporation. 0 0

Reduction of heat losses by appropiate insulation of equipment and unit parts Solution heating by vapor and vapor condensate of the evaporation unit Multistage solution evaporation including vapor heat use (the specific steam consumption dH decreases with increasing number of stages n proportional l/n) Estimated data of specific steam consumption of single o r multistage evaporation of aqueous solutions [7.22] Number of stages

2

1

4

6

Specific steam consumption without thermal compresssion (t/t, related to steam and solution vapor)

1.1 (1.0)

0.6 (0.5)

0.3 (0.25)

0.2 (0.17)

Specific steam consumption with thermal compression (t/t, related to steam and solution vapor)

0.5

0.33

0.17

0.12

Values in brackets are valid for ideal evaporation units 0

0

Single or favored multistage solution evaporation including vapor compression by mechanical or thermal means (additional steam savings by energy supply to vapor, which are usually condensed and fed back to the evaporator) Multistage flash evaporation Use of waste heat by heat transformation (see Chapter 2.5.2.3)

C, ( n )

uln c ln

u

0

I I

No. of stages n

Co(n)

nopt __F

Fig. 7-11.Optimal number of stages of a multiple effect evaporation unit.

Figure 7-12 shows different variations of solvent flow in a continuously operated multistage evaporator unit. With cocurrent flow which is usual in practice (Fig. 7-12a), the solution flows parallel to the vapor through the evaporation sections. The concentrated solution is treated in the final stage at the lowest operating pressure, which may lead to poor heat transfer in a highly viscous solution. Transfer pumps between the individual stages are generally not required if the stages are operated at different pressures. In countercurrent flow (Fig. 7-12b) the solution and steam flow countercurrently. Dilute solution enters the last stage at the lowest operating pressure, is

7.2 Crystallization from a Solution IG

cw

489

Fig. 7-12. Different methods to connect evaporators in a multiple effect evaporation unit, demonstrated with a two stage evaporation unit.

a) Cocurrent (forward feed)

b) Countercurrent (backward feed)

c) Parallel connection

EV Evaporator BC Barometric condenser TP Condenser tail pipe F Feed, dilute solution cs Concentrated solution SV Solvent vapor so Solvent IG ST Steam Condensate .cw co IG Inert gas cw Cooling water EF Effluent (waste water)

bl

ST+

CS C)

r----

57

4 cs

each evaporator and a final concentration is achieved in each unit. Steam is guided through the evaporation section in the same manner as in the two previous variations. Individual arrangements are often combined to obtain optimum operating conditions and suitability to the solubility behavior. In multistage operation, it is assumed that each evaporation section operates with a small difference between the steam temperature and the solution boiling point, and that they operate economically. A procedure to calculate the steam requirement DH for a continuously operatedstage evaporation unit in parallel flow, including preheating of the solution by some of the evaporated solution, is given in Table 7-3. Connections, flow sheets and nomenclature can be found in Fig. 7-13. Further heat saving for the process according to Fig. 7-13, is achieved by using the condensation heat of the vapor to preheat the solution. In this way, some of the vapor generated by the evaporation, may be withdrawn to be used elsewhere [7.20]. The condensation heat of the vapor may also be used as waste heat for heat transformation .12.53., 7.231.

ITP

concentrated from stage to stage and, as a final concentrate, leaves the first stage at the highest operating pressure. Heat transfer problems due to high viscosity of the solution are not observed, but the disadvantage is that the solution has to be pumped from stage to stage. In a parallel arrangement (Fig. 7-12cjdilute solution is fed to

7 Solvent Evaporation, Crystallization

490

Table 7-3. Procedure to calculate the steam consumption DH of an n-stage cocurrent evaporation unit including solution preheater (connections, flow sheet and nomenclature may be found in Fig. 7-13).

Equation system: 0

Balance area: Stage 1 (7-9) Balance area: Stage I and Stage 2

(7-10)

(7-11)

(7-12) 0

Total evaporated solvent: D=Fa.(l-!E)=D,+D

+..

+ Dn = f(&)

(7-13)

Procedure to calculate DH 0 0 0

0

Selection of pressure and temperature levels Determination of all enthalpies Calculation of partial solvent vapor flow rates for preheating purposes Dv, - Dv,n by Eq. (7-12) Calculation of solvent vapor flow rates D, - D, by Eqs. (7-9)-(7-11) Calculation of steam consumption by Eq. (7-2)

The heat exchange area A i for the evaporator stage i is Qi

= ki * Ai * Adj

(7-14)

where Qi is the heat exchanged in the evaporator stage i, ki is the heat transfer coefficient and Adi is the effective temperature

difference. The total available temperature difference A d , is determined by the steam temperature dH and the vapor temperature dB,n leaving the last stage n of the evaporator

Ad, = % - dB,n

(7-15)

7.2 Crystallization from a Solution

.

.

I .... +Solution ---,-- Solvent vapor +Solvent

Stage n

491

-1.E"

....

-Steam --+--Condensate

Fig. 7-13. Steam rate of a multistage evaporation unit. EV Evaporator PH Preheater

Note that this temperature difference is not completely available for apportioning into individual evaporation stage temperature differences ALP], ALP2, ..., ALPi, ..., ALPn. This is due to the raising of the boiling point A T in each stage i and heat losses in the vapor pipe system which cause the vapor to cool by ca. 1-1.5 K. These temperature losses in each evaporation stage increase with increasing solvent concentration and increasing number of stages, resulting in a total temperature loss ALPu for the entire unit

cAT + n

ALP" =

1

(1.0 ...1.5) * (n - 1)

has to be divided in such a way that each vaporizer has the same heating area with respect to the lowest manufacturing, maintenance and service costs. With the valid proportionality

... : ALP,

ALP, : ALP2 :

81 . Q2 . ... . Qn

--

kl

*

k2

* -

kn

=

(7-18)

the effective temperature difference ALPj for any stage j of the evaporation unit follows

(7-16) (7-19)

The remaining usable temperature difference n

C ALP, = ALP, - ALPu 1

(7-17)

This calculation assumes knowledge of the heat flows Qi and the heat transfer coefficient ki. Since Q, and k, are generally not

492

7 Solvent Evaporation, Crystallization

known in practical design for an evaporation unit, temperature difference Adj are assumed for iterative calculation [0.6], [7.20]. To determine the heat transfer coefficient ki in Eq. (7-14), see, for example [0.1, 7.24-7.261 covering the field of heat transfer problems in evaporation processes. During solution evaporation and crystallization, fouling and incrustation of the heat exchange areas has to be considered, which may lead to substantially lower evaporation rates in the evaporator. With incrustation on the product side of the evaporator tube wall, the heat transfer resistance is 1

1 6, -=-+-+ -+kef, “i A,.

6,

1

A,

a,

(7-20)

where a , a, are the heat transfer coefficients of the product and heating-medium sides, ,a, Afi are the thickness and heat conduction of the fouling layer and ,a, A, are the thickness and heat conduction of the tube wall. Therefore, the so-called fouling factor, R, is (7-21) where k , keff are the heat transfer coefficient for the clean and “fouled” heat exchange area. In [7.27] and [7.28] a detailed discussion is given about what causes fouling, the affects of fouling and the associated mechanisms. Determination of the fouling factor for certain operating times of the evaporator until the next service cleaning are only based on empirical data. 7.2.1.3 Solution Evaporation with Mechanical and Thermal Vapor Compression

At the operating pressure pA a part or the entire solvent vapor from the evaporator, is

compressed to a higher pressure p N using the principle of an open heat pump cycle, and may then be used as heating steam. Vapor compression results in a saving of external steam flow to the unit, and is applicable in single as well in multistage evaporation units. Vapor compression is carried out by mechanical compression (reciprocating compressor, lobe compressor, axial-flow or radial-flow compressor) with an electric motor or turbine as a propulsion system, or thermally by means of a vapor-jet compressor. With mechanical vapor compression (Fig. 7-14; see also Table 2-12) the vapor from the solution is removed from the vapor space of the evaporation unit slightly superheated due to the rise in boiling point (1 -+ 2), polytropically compressed (2 3) and then recycled to the heating section of the evaporator, where heat is withdrawn under isobaric conditions (3 + 4 5 ) (Fig. 7-14c). The effectiveness number Eeff is the ratio of saved or generated effective heat and the required compression energy. With a decreasing pressure difference p N - pA or pressure ratio pn/pA, Eeff increases, but a decreasing pressure difference or pressure ratio implies a smaller usable temperature difference in the evaporator and therefore requires a larger heat exchange area. If the energy and investment costs of the evaporation unit are plotted against the pressure ratio pN/pAand both curves are superimposed, an optimal value is observed at the minimum of the total costs curve. Generally, mechanical vapor compression is of economic sense if the achieved vapor temperature difference, up to 20 K, is enough to operate the evaporator. To determine the most cost efficient design for an evaporation unit, the energy cost per ton evaporated vapor has to be evaluated for each connected variation. A nomogram to find the energy cost for each ton of water evaporated from an aqueous solution -+

-+

493

7.2 Crystallization from a Solution a)

-Start-up RV steam

Surplus

--Surplus vapor Heat losses --Concentrated solution

I(

--

*- Dilute solution I ondenzte

C o n c e n t r a t e d Dilute solution solution

Condensate

PN/

5-

s-

Fig. 7-14.Evaporator with mechanical vapor recompression. a) Flow sheet b) Energy flow diagram is given in Fig. 7-15. In the left part of the c) Temperature/entropy diagram figure, the specific energy consumption for d) Enthalpy/entropy diagram mechanical vapor compression is presented EV Evaporator as a function of the vapor temperature or CO Compressor vapor pressure and the temperature differSV Solvent vapor ence AT in the compressor. On the right RV Recompressed vapor PA, Phi Pressure before and after vapor side, the energy cost per ton evaporated water is plotted as a function of the number of compression ATs Boiling point elevation stages of a multistage evaporation unit LC Limit curve against the costs of heating steam. A link

between the diagrams is possible via the current cost line. Application of the diagram may be explained by an example (different energy costs require adjustment of the diagram to the appropiate currency): A solution with a rise in boiling point of 8 K has to be evaporated. The steam price is 25 DM/ton, and

7 Solvent Evaporation, Crystallization Vapor temperature l°Cl

- +

. .

~

60

'fi

70

90

80

I

100

1

:

ar = 10 K AT=l?K

ATrlHK

AT

123

~

70 K

1W

--

312

473

101

1013

Vapor pressure lrnborl

a

io

20

-

30

40

'$1

&

S!eam price IDM/+I

Fig. 7-15. Nomogram to find the specific energy costs for each ton of water evaporated from an aqueous solution for the case of mechanical vapor compression and multistage evaporation. Representation according to Mannesmann-Engineering AG, Messo-Chemietechnik [7.29].

the electricity cost is 0.15 DM/kWh. The boiling point of the solution is 9 3 ° C and by considering the rise in boiling point, the vapor temperature at the compressor inlet is 85 "C. With a sufficient temperature gradient of 7 K on the heat exchange area, a ternperature rise of the vapor of 7 + 8 = 15 K follows and is accomplished by pressure increase caused by vapor compression. The procedure may be found from the dashed point line in the nomogram. For mechanical vapor compression, 5.50 DM/ton evaporated water has to be paid out for electrical power costs. Steam prices for multistage units operating with the same energy costs

as a single unit including vapor compression are 15 DM/ton for three stages, 16.50 DM/ton for four stages, 21 DM/ton for five stages and 25 DM/ton for the six stage unit. With the chosen steam price of 25 DM/ton, steam costs are 10.40 DM/ton evaporated water for the three stages to 5.50 DM/ton for the six stage unit [7.29]. With thermal vapor compression (Fig. 7-16) by means of a vapor jet compressor, the steam energy at a higher pressure is used to raise the vapor, generated by evaporation, to a sufficient pressure level, such that the compressed vapor may be used as a heating medium.

7.2 Crystallization from a Solution

495

bl High pressure steam

-

-

Residual solvent vapor

Residual solvent vapor

Heat loss

Concentrated solution

-Dilute solution

High pressure steam

I

i

Condensate

Concentrated solution

Fig. 7-16. Falling-film evaporator with steam jet compressor (a) and energy flow diagram (b).

Representation according to GEA-Wiegand, Ettlingen. FFE Falling-film evaporator SC Steam jet compressor SS Suction solvent vapor HS Heating steam

Figure 7-17 explains the function of a steam jet compressor. Inside the booster nozzle, frequently named Lava1 nozzle, designed with a diverging cross-sectional area followed by a converging cross-sectional area, the booster steam is relaxed (1 2 or 1 -+ 2'). Here, pressure energy is converted into kinetic energy. Booster steam is mixed with the extracted vapor in the entry cone of the mixing nozzle (2 3 or 2'+ 3, state point 4 or 4'). By momentum exchange, energy is transferred to the vapor, and the vapor is also accelerated. The kinetic energy of the obtained mixed vapor or steam is converted to pressure energy by means of a pressure jump or compression shock, and by de-+

+

lay in the diffuser through compression to the exit pressure (counter pressure) p N (4 -+ 5 or 4' 5'). The achieved counter pressure p N lies between the booster pressure p r and the intake pressure p A of the vapor. For the steam flow rates DT (booster steam), bK(intake vapor) and Dn (mixed vapor or steam) the required heating steam for the evaporation stage is, according to Eq. (7-2) and using the nomenclature of Fig. (7-16), -+

DH = d, + DK

(7-22)

With the conveying ratio E for the steam jet compressor

496

7 Solvent Evaporation, Crystallization

I

b) BS

t-'K.

I

Suction vapor PA

I F l o w direction

I

C)

Mixed vapor. -)DH.PN

-

Pl

h

PA

I

L

s-

Fig. 7-17. Steam jet compressor. Representation according to GEA-Wiegand, Ettlingen. a) Construction b) Pressure curve c) Enthalpy/entropy diagram 1 . . 2 ; 1 . .2' isentrope, i. e., polytrope (subject to loss) removal of booster steam 2 . . 3 -+ 4 or 2'. . 3 -+4' mixture of booster steam and suction vapor 4 . . 5 ; 4'. .5 ' isentrope, i. e., polytrope compression of mixed vapor 1 Body 2 Booster nozzle 3 Inlet cone 4 Neck 5 Outlet cone (3, 4, 5 diffusor, mixing nozzle) BS Booster steam SV Suction vapor MV Mixed vapor LC Limit curve

7.2 Crystallization from a Solution

DK

&=Y

The partial vapor flow rate f i K to be compressed is

(7-23)

DT

or its reciprocal value, the specific booster steam required, d,/&, the required booster steam flow rate D T follows from Eq. (7-22) DT=DH.---

1

fiK -- fiH ' -

&

&+l

(7-25)

Figure 7-18 shows the specific steam consumption 1/& as a function of the vapor intake temperature rSA or vapor intake pressure pA, and the required temperature in-

(7-24)

&+l

497

20

.

10

1

a

[kglkgl 1

2

1.o 0.8

0.6 0.L

0.2

0.1

20

,

.+

LO

60

'

I

-

80 ' 100 S,[OCI

'

I

I

1 i O ' 1l'O

'

Id0

Fig. 7-18. Specific steam consumption of steam jet compressors with a propellant steam rate of about 100 kg/h. Representation according to GEA-Wiegand, Ettlingen. D,/D, Specific steam consumption tpA Vapor suction temperature _ _ - pH= 4 bar absolute _ . _ . p H = 10 bar absolute p H = 20 bar absolute ~

498

7 Solvent Evaporation, Crystallization

crease $ !I, - dA, by compression. With increasing booster steam pressurehtake pressure and decreasing required temperature rise, the booster steam consumption decreases. Therefore, thermal vapor compression is of special interest with low intake pressures, i. e., vacuum operation of evaporators, and small compression ratios p N /p A .

7.2.1.4 Multistage Flash Evaporation

Concentration of a solution is also possible by multistage expansion evaporation [7.21, 7.301. The solution is preheated in the stages of the evaporation unit by the flashed solvent vapors and externally supplied steam in a final preheater to a boiling temperature T, at an elevated pressure. Then it is expanded from stage to stage and flash evaporation occurs. During the expansion process, the solution is concentrated by separating the generated vapor. Multistage expansion evaporation is characterized by low usage of external heat. Another advantage is, that the actual evaporation does not take place on the heat exchange area and therefore no incrustation of the heating area occurs.

A scheme for an n-stage expansion evaporation unit with usage of the vapor heat is presented in Fig. 7-19. It can be used, for example, for the desalination of sea water. If a boiling solution of mass flux, i,is partially expanded in a constriction by experiencing a differential pressure drop dp, the solution is cooled by d T and expansion is generated (expanvapor of flow rate sion evaporation under isenthalpic throttling at boiling conditions). A heat balance over the constriction with steam separator gives

dd

Ah[,, = L . cP,L . d T

(7-26)

if no further crystallization in the solution occurs, and the change of temperature is only due to a change in enthalpy of the solution. The first stage of the expansion evaporation unit in Fig. 7-19 gives the flashed vapor flow rate D,

when the entering solvent flow L , is cooled by AT, caused by an expansion of A p , . cpLt and Ah,,gl are the mean specific heat of the

Con d e n sate Concentrated

Fig. 7-19.Multistage expansion evaporation unit, including solution preheating by vapor. FP Final preheater SP Solution preheater CE Condensate expansion SE Solution expansion SV Solvent vapor

7.2 Crystallization from a Solution

solution and the evaporation enthalpy of the solvent in the temperature range t9, CS2. D, is only considerable at high rates of L , and/or extended throttle effects. This is also naturally true, for other flashed flows and also for the total vapor flow D of the expansion stages -+

D = C1 D , = L , . ( l

-?)

or alternatively dS =

- A X - dL

1 + Ax + (1 - p )

. Ax

=b.& (7-33)

Integration of Eq. (7-29) gives, for example, for a single expansion stage

(7-28) L2 In-;-=

The release of vapor causes an increase in the concentration of the solution from an initial mass fraction w, to a final mass fraction w,, in the final stage. Expansion evaporation is only of economic sense when a small concentrating effect w , w, is sufficient or when evaporation of a small amount ensures enough solution supersaturation, for example, in vacuum crystallization. If not only solvent, di,is evaporated but also solid dS crystallizes during the expansion, then the balance in Eq. (7-26) becomes

499

L,

T2

-

j

TI

CpLl

a.

+

dT (7-34) *

hCl

The heat consumption QH for the final heater (FP) in the expansion evaporation unit is

-+

(7-35)

if Tvmis the temperature of the solution after the final vapor heated preheater (SP) and T, as the entry temperature into the first expansion stage (SE) (Fig. 7-19). With the assumption of ATE equal for every preheating stage (SP) and every expansion stage (SE) (7-36)

where h, is the crystallization enthalpy, referred to the crystal. The change of solvent flow d i during the expansion is dL = dB + dS + (1 -,u). dS

(7-30)

where p is the solid fraction of the crystals. It then follows with A X as the crystal solid mass per unit mass of evaporated solvent, dS=AX.dd

and including the relationships of dS to dL

(7-31)

and

where AT, is the achieved temperature difference in the expansion, AT is the rise in boiling point and A T the required minimum temperature difference between the last preheater and the final heater. The specific heat consumption q for the final heater and therefore the specific head consumption supply QH q=-=

D

LI

c p - ~(TI -

D

Tvd

(7-37)

decreases with increasing number n of expansion stages, with increasing preheating -& temperature T, and with decreasing differ=a.dL dD = ence between the vapor temperature and the 1+ A x + ( l - p ) * A x (7-32) solvent exit temperature of the preheaters

500

7 Solvent Evaporation, Crystallization

(SP). Data for desalination of seawater, for example, are given in [7.31, 7.321. Expansion evaporation may also be applied in a multistage evaporation unit to concentrate a solution from stage to stage at lower operating pressure, by throttling before each following stage. For the practical implementation of expansion evaporation, a final heater and a series of vaporAiquid separators and preheaters, including throttle valves, is sufficient. Modern multistage expansion evaporation units are compactly designed to reduce heat losses and space requirements.

7.2.1.5 Types of Evaporators to Concentrate Solutions

For solution evaporation, circulation evaporators are often used in the form of a tube bundle apparatus with vertical, horizontal or slanting tubes, operated with natural or forced solution circulation. Tubes may be mounted inside or, with respect to better cleaning, outside (Fig. 7-20) of the evaporator. The solution evaporates inside the tubes and is lifted by the vapor bubbles generated according to the principle of an air lift pump. Natural circulation of the solution is mainly intensified by adding circulation pumps (forced circulation). Circulation evaporators are characterized by large liquid holdups, large liquid loads and good heat exchange, but also by larger mean residence times. A high pressure drop appears, especially with forced circulation in narrow tubes. Thermally gentle treatment of a solution is enabled by continuous (one pass) evaporators (Fig. 7-21). The solution flows as a film through the evaporation zone, driven by gravitational force, and may be additionally whipped, or dragged up by the vapor. A small pressure drop, low liquid holdup and hence a small mean solution residence

time distribution, are advantages of these evaporators. Discontinuous solution concentration is carried out in heated stills, similar to these used in simple discontinuous distillation (Fig. 2-6), and evaporators with stirring gear (Fig. 7-22). A brief overview of selected evaporator designs including considerations for design and operation is given in Table 7-4.

7.2.2 Balancing of Crystallizers The crystallization yield (crystal product rate) and supplied or removed heat flow due to crystal formation in the solution, result from mass and energy balances over the crystallizer, in single or multistage crystallization plants.

7.2.2.1 Crystal Product Rate The crystal product rate expected during the crystallization process may be calculated from a mass balance over the crystallizer with known feed and discharge flows and the substance properties of the dissolved substance in the solvent, such as solubility behavior, crystal load, etc. A concentrated solution of flow rate with a saturation load X , , at entry temperature C$, given in kg dissolved substance per kg solvent, is fed to the crystallizer (Fig. 7-23). (4 is identical to F, in Chapters 7.2.1.1 and 7.2.1.2, if the solution is preevaporated to the saturation point. X , therefore can be found from w,

6

x,= 1 -w,w, ~

(7-38)

if 19, is the exit temperature of the solution from the last evaporation stage and is also

Vapor

-1

f

Vapor

e

I \'apo-

t

I

~

1

d d

'IL

I

Feed

t

Concentrate

Concentrate -Vapor

b

Fig. 7-20. Different types of circulation evaporators. [0.1, Vol. 21 and data of GEA-Wiegand, Ettlingen. Representation according to SCHAEFER I Robert thermosiphon evaporator I1 Herbert slanting tube evaporator 111 Forced circulation evaporator IV Three chamber circulation evaporator V Forced circulation evaporator, horizontal position a Evaporation section b Downcomer c Separator d Heating agent e Vapor section f Circulation pump g Concentrated solution pump

502

7 Solvent Evaporation, Crystallization a)

--

b)

Vapor outlet

!

Cl

Vapor outlet

1 I

"S

VS

SE

Vaporoutlet

SE

solution inlet

3

Concentrated solution outlet

I

f- Dilute

)

ES ES RO

ES

3-E 1

D&

ES

4-E LRB solution inlet Concentrated solution outlet

I

t

DR

Concentrated solution outlet

dl

RIB RFB

solution inlet

7.2 Crystallization from a Solution

503

C)

so

fi-Solvent

va

vapor

Dilute-

Dilute-!

solution

HA HA

1 I Concentrated

solution

L , HA

-

+--HA

solution HA-

/\

I

Concentrated HA

Concentrated

solution

solution

Fig. 7-22. Evaporator with stirring gear, discontinuous solution evaporation. Representation according to GEA-Wiegand, Ettlingen. a) Evaporator with propeller stirrer for solutions of medium viscosity and good flow behavior b) Evaporator with anchor stirrer for solutions of high viscosity c) Evaporator with paddle stirrer and heated hemispherical head for solutions of very high viscosity and unfavorable flow behavior HA Heating agent

4 Fig. 7-21. Different types of continuous

(one-pass) evaporators a) Falling-film evaporator *, countercurrent flow of liquid phase and vapor phase b) Climbing film evaporator * c) Thin-film evaporator ** d) Rotor types** Representation according to * GEA-Wiegand, Ettlingen. ** Luwa-SMS Verfahrenstechnik, Zurich/ Butzbach.

I ES SE VS DR RO URB LRB RIB RFB D E

Inert gas inlet Evaporation section Separator Vapor section Drain Rotor Upper rotor bearing Lower rotor bearing Rotor with inflexible blades Rotor with flexible blades Steam inlet Condensate outlet

504

7 Solvent Evaporation, Crystallization

Table 7-4. Selected evaporator types including design, operation and references. Agitated evaporator (Fig. 7-22): Discontinuous evaporation of solutions with high viscosity or even pasty or pulpy consistency. Heated by 0 Double-walled jacket Low evaporation rate due to the small ratio of heating area A to content Vand decreases with increasing vessel diameter d, ( A N - l / d ) , heat transfer coefficient k = 700-1700 W/(m2. k) [7.33] 0 Welded tubes or semicircular tubes, low evaporation rate, k = 700-2300 W(/m2. k) [7.33] 0 Internal heating coil (for nonincrusting products), k = 1200-3500 W/(m2 . k) 0 External heating with pump circulation 0 Direct firing 0 Electric heating jacket (laboratory evaporator)

Heated stills Internally heated with horizontal boiling tubes (submerged-tube evaporator), continuous or discontinuous operation, k = 1000-3000 W/(m2 . k) at water evaporation and under a vapor pressure of 1-7 bar [7.21], heat transfer calculations according to [7.34]. Circulation evaporator (Fig. 7-20): Continuous evaporation of solution flowing upward in tubes (vertical or slanted mounted tubes) or from a solution passed through shell side (horizontal mounted heating tubes). Reflux of excess solution after vapor separation into the evaporator (circulation). 0 Natural-circulation evaporator with short tubes (Robert thermosiphon evaporator) : Circulation flow of solution according the thermosiphon principle, i. e., density differences of solvent vapodsolution in evaporator tubes and solution in the central downcomer, boiling tube diameter d = 25-70 mm, boiling tube length Z = 800-2500 mm, heating area A = 25-350 m2, evaporator diameter d,,, = 1500-2500 mm, downcomer diameter: ca. 200-800 mm, downcomer cross section/total cross section of all boiling tubes: ca. 0.5-1 [0.6], k = 350-900 W/(m2 . k) (viscous solution), k = 600- 1800 W/(m2 . k) (dilute solution) [7.21], calculation of heat transfer according to [7.35], solution residence time up to 30 minutes and longer. Inclined tube evaporator: The slanted mounting of the tubes leads to a lower hydrostatic liquid height (hydrostatic pressure), therefore, to a lower liquid content with respect to the evaporation rate and a higher heat transfer than the Robert evaporator [7.21]. Forced circulation evaporator with external heating : Forced circulation of the solution is guaranteed by a circulation pump and leads to an increase in heat transfer at lower difference between wall temperature and boiling point. Usually the evaporation section is shifted into the vapor space to reduce incrustation of the heating area, cleaning and maintenance are simpler than with the Robert evaporator, Z = 1500-4000 mm, d = 20-60 mm, Z/d = 30-60, A = 100-900 m2, d , = 1500-2500 mm [0.6], k = 900-3000 W/(m2 . k), to calculate the heat transfer see [7.21].

7.2 Crystallization from a Solution

505

Table 7-4. (continued)

Continuous (one-pass) evaporator (Fig. 7-21): Continuous evaporation from a rising or falling film in the evaporator. One-pass, low hold-up, low mean residence time of solution in the evaporation section, due to low vapor pressure drop more gentle evaporation under vacuum operation than the circulation evaporator. 0 Falling-film evaporator : At the top distributed solutions flow down the inside of evaporator tubes as a falling film by gravitational force, vapor is evaporated from the liquid film, Z = 5-7 m, d approximately 50 mm, liquid load ca. 0.7-3 m3/(h. m), mean residence time: 1-3 min, to calculate the heat transfer see [7.21]. 0 Thin-film evaporator with rotating stiff or swinging wipers: At the top feed liquid flows downward as a thin film by gravitational force, by means of a highspeed rotor film thickness is equalized and the surface is constantly renewed. Z = 0.5-0.6 m (heated length), d = 0.15-1.7 m (outer diameter), heating area A = 0.1-25 m2, operating pressure down to ca 1 mbar, mean solution residence time 10-60 s, rotor speed ca. 1000 revolutions/min, film thickness 0.1 - 1 mm, final solution viscosity ca. 20-200 Pa . s, k = 500-1800 W/(m2 . k) at concentration of aqueous solutions, evaporation rate ca. 80- 1000 kg/h water, to calculate heat transfer see [7.21] Thin film evaporator with rotating radiator: By the effect of centrifugal forces the solution is spread as a thin film (mean ca. 0.1 mm) over a conical frustum, rotating with 1600 revolutions/min. Short residence time of ca. 1-2 seconds, relatively low wall temperature on the product side, treatment of higher viscosity, temperature-sensitive products, high heat transfer coefficient of 500- 1000 W/ (m2 * k) depending on evaporator body revolutions, product and operating conditions [7.21]. 0 Long-tube evaporator: Solution flows with high velocity through the long evaporation tubes from bottom to top whilst solvent evaporates. In the top section of the tubes liquid is dragged along as a thin film by the vapor, thus a fully developed two phase flow with a good heat transfer is obtained. 2 = 4-9 m, d = 20- ca. 50 mm, A = 100-900 m2, d,,,, = 800-2000 mm [0.6], to calculate the heat transfer see [0.17]. Note: Overall heat transfer coefficients refer to condensing steam as a heating agent.

from d, to 4 thereby becomes supersaturated. Solvent is simultaneously evaporated at the vapor flow rate &. Reducing supersaturation during the crystallization process and subsequent phase separation, a crystalline phase S and the remaining saturated so(7-39) lution at temperature d2 with a mass ratio X,, are obtained. Introducing the ratio

the entry temperature of the solution to the crystallizer). The fraction of pure solvent in the feed solution is L,,

=

6

~

1+

x,

To determine the crystal production rate S for solvated crystals the general case of combined cooling and evaporation crystallization is postulated. A solution cooled

(7-40)

506

7 Solvent Evaporation, Crystallization = o), pure evaporation crystallization (bK zation (XI = X , = X , at the evaporation temperature dS) and L,I = 1 for the formation of salt free from crystal liquid. The yield q of the crystallization process is proportional to the crystal production rate and is given by

(7-43)

Fig. 7-23. Crystal product rate and heat rate. CS Concentrated solution ML Mother liquor VA Vapor CR Crystal

of the molar mass M , of unsolvated crystals and the molar mass M k of solvated crystals, a mass balance for the salt gives

i ,. x,= 8 . p + [iTl -& - s . (1 -/A)] . x, (7-41) where L , XI salt in the feed solution 8-p crystals or salt free from crystal liquid S(1 - p ) solvent fraction of solvate [ L T l - & - s . (1-p)] salt in the remaining solution +

e x ,

Rearranging Eq. (7-41) with respect to the crystal product rate, 8 gives

From Eq. (7-42) special cases are easily deduced, for example, pure cooling crystalli-

Since the remaining solution leaving the crystallizer is usually still supersaturated, the calculated yield in Eq. (7-43) is not reached in practice, Also, reliable solubility data to calculate are seldom available for an exact calculation of S . If two substances totally crystallize in a solution, three mass balance equations are formulated. Saturation isotherms, crystallization path, crystal yield and crystal composition can then be presented in a triangular diagram and determined [0.26].

7.2.2.2 Heat Exchange During Crystallization

The heat flow Q supplied or removed during crystallization is derived from a heat balance over the crystallizer (Fig. 7-23). For the general case of a combination of a cooling and evaporation crystallization,

F1 * h , + Q = = D K . hD,K

+ F2. h2 + S * h, + Qc + Qu

(7-44)

where

Q is the supplied heat flow F, . h , ,F2. h, heat of the feed and heat of the remaining solution leaving the crystallizer with enthalpy h l , h2

7.2 Crystallization from a Solution

(7-45) (7-46) and the specific heat c4,1 , the feed and remaining leaving solution

cp,2 of

removed heat flow rate carried by the evaporated solvent enthalpy of the solvent vapor (to be calculated analogously to Eq. (7-4)) heat removed by the crystal phase

507

Special cases of pure cooling crystallization (& * h , , = 0) and pure evaporation crystallization can be easily derived from Eq. (7-44). Energy dissipation by stirring or circulation devices is not considered. The heat flow Q is exchanged with the crystallizer by means of internal or external heat exchangers, of area A

Q

A= ~

E

*

(7-50)

E

where k,,, is the effective heat transfer coefficient, calculated analogously to Eq. (7-20), if the fouling factor R, and the (7-47) outer and inner heat transfer coefficients are known and with Aff, as the logarithmic mean temperature difference between the hs is the enthalpy, and heating or cooling medium and the circuthe specific heat of the cryslated solution or suspension. To provide fatal phase vorable conditions for coarse crystal formaheat effect during crystallition, and to avoid spontaneous crystal seed zation formation in the relatively nonmoving - exothermic crystalliboundary layers of the heating tube walls, zation AW, or supersaturation of the solution + endothermic crystalliAX, should be smaller than, AW,, or zation (7-48) AX,,, of metastable solutions (distance between solubility curve and second supercrystallization enthalpy de- saturation curve, Fig. 1-41). For a cooling of pending on temperature, Ad, in cooling crystallization, loading, and type of solvate heat losses from the crystallizer The crystallization enthalpy hc has to be found experimentally. An approximation is possible according to Chapter 1.4.6 if points q,XI and q,X , on the solubility curve are known -

h, =

R Tl . T2 (InX2 - InX,) 7-2 -

T,

(7-49)

The correct sign of Qc should be used in application of Eq. (7-44). Q > 0 indicates heat supply, Q < 0 indicates heat removal.

= Aff

-

(g)i

AW,,

(7-51)

where dw/dd is the solubility gradient (slope of the solubility curve) and csL is the specific heat of the solution. To avoid or to inhibit incrustation of the heat exchangers with cooling crystallization, the 0

Tube wall surface should be as smooth as possible, coated if necessary (PTFE)

508 0

0

0

7 Solvent Evaporation, Crystallization

Difference between the tube wall temperature and solution temperature at the cooler exit has to be adjusted so that it is smaller than the temperature difference of a metastable solutions (Ad,,, Fig. 7-8), i. e., sometimes only a few Kelvin depending on the substance system Velocity in the tubes should be as large as possible, if needed secondary unwanted seed formation has to be considered Cooling rate has to be small and controlled t0.26, 7.361

0

0

0

With evaporation crystallization, the supersaturation Aw, or AX, is controlled by the evaporation rate DK/Fl

where ws is the mass fraction of the dissolved substance in the just saturated solution. To avoid incrustation in the heating devices of the evaporation crystallizer, the following must be considered: 0

0

0

0

The evaporation rate has to be restricted and controlled The solution heating in the evaporator may only be so high that the resulting evaporation per cycle does not cause a large supersaturation Aw,; temperature differences between the tube wall and solution are ca. 10-20 K The residence time of the supersaturated solution should be at most 10% of the residence time in the crystallizer Evaporator type and dimensions have to be chosen such that sedimentation of dissolved substance does not occur anywhere. The diameter and number of the boiling tubes (min. 25 mm) have to be set to ensure a solution velocity of 1.5-2 m/s, whilst still avoiding secondary seed formation (see Chapter 7.2.3)

Forced circulation eases the continuous throughflow through all suspension areas and guarantees constant velocity and temperature differences The distance between the top heating tube bend and the solution level must be larger than the corresponding vapor pressure difference between the heated and expanded solution. This is to avoid blanching in the heating area and, therefore, spontaneous incrustation (estimated data ca. 200-300 mm) The vapor space should be sufficiently large to avoid carry over of solution by the vapor. The vapor velocity w v should be smaller than 1.5 times the vapor bubble rise velocity, given by [7.37]

where e, and egare the density of the solution and the vapor, respectively and is the surface tension. During solution cooling and evaporation, it must be noted that the state point of the mixture of the fresh and circulated solutions lie in the metastable solution range. The mixture fed to the heat exchanger should contain a sufficient amount of crystals (suspension density =3-10%), and so that a reduction of the supersaturation occurs by growth of the circulated crystals and not by growth on the tube walls. Addition of additives may also avoid incrustation.

7.2.3 Crystallization Kinetics, Crystal Seed Formation, Crystal Growth Crystallization actually occurs by reducing the supersaturation of a solution, which is achieved by cooling, evaporation of solvent

7.2 Crystallization from a Solution

or by combined cooling and evaporation under vacuum (Fig. 7-8). Apart from specialties, such as fast crystallization or onecrystal growth, two stages are required to reduce supersaturation. In the first stage crystal seeds are formed (nucleation), whilst in the second stage, these crystal seeds grow, and exceed a critical minimum size by taking up solid out of the supersaturated solution to form coarse crystals. To obtain a product which is as coarse as possible, the supersaturation reduction should take place in the metastable area between the solubility curve and the second supersaturation curve (see Chapters 1.4.5.1 and 7.2.2). The number of crystal seeds formed per unit volume of the solution and per unit time, the nucleation rate, mainly depends on the degree of supersaturation of the solvent, AX,

AX,=X-X,

(7-54)

where X i s the loading of the supersaturated solution and X , is the equilibrium load, or saturation load, as given by the solubility curve at the reference temperature. Other factors also change the nucleation rate and at the same time, control the width of the metastable area. These include solution impurities, pH value, viscosity, flowing state and temperature of the solution. If the seed formation is primary and statistical, i. e., with homogeneous seed formation in the solution, the nucleation rate is given by an Arrhenius equation of the form

dt

'

where C , and C, are constants depending on the system. With the heterogeneous formation of seeds, the crystallizing substance bonds it-

509

self to the crystallizer wall or foreign particles (e. g., impurities). In technical crystallization processes, primary (homogeneous and heterogeneous) nucleation is unimportant. In this case, seeds are mainly formed by the collision of two crystals with each other or with the crystallizer wall, and are also formed by shear stress, particularly in the areas influenced by stirring devices and circulating pumps. Methods to describe the secondary nucleation rate are given in [7.38, 7.391. For example, an empirical approach for the effective nucleation rate B, as the number of seeds per m3 suspension and time in seconds is

where k is a system constant, er is the suspension density (mass of crystals per unit volume kg/m3) and E is the dissipated energy per unit mass in W/kg ( E = 0.1 - 1 W/kg in agitated crystallizers). The linear crystal growth rate rg is given by dx r =dt

(7-57)

where x is the crystal diameter. The exponents a, b, c are found by experimentation (a = 1, b = 0.5 ..-1, c = 1.3 .-2.3) for different aqueous solutions as described in [7.39]. The nucleation rate increases with increasing supersaturation; spontaneous formation of many small nuclei is observed particularly after passing the second supersolubility curve (see Fig. 7-24). The methods required to control nucleation in a crystallizer include Avoiding supersaturation in the area of an unstable solution Limiting the crystal growth rate Reducing mechanical stresses on the crystals during agitation and circulation by pumping

510

7 Solvent Evaporation, Crystallization

1

In the case of a continuous degree of supersaturation, crystal growth is described by [7.43] 1 rg

-

1

r,

+ C,

1

w,

+-

ri

(7-58)

where rg r,, r,

C,

w, Fig. 7-24. Nucleation rate, crystal growth rate, and crystal growth ratehucleation rate ratio as functions of supersaturation.

Pretreating the solution, if necessary, by adjusting the pH value and adding additives Discharging the fines fraction separately, and if necessary, granulating or dissolving the fines (classifying crystal discharge and dissolving of fines) [7.38] Seed crystals added to the solution and/or the seeds formed in the solution grow to larger crystals as long as the solution is supersaturated. Crystal growfh is a complex process, which is assumed to consist of two steps. During the first step, solid molecules diffuse from the solution bulk to the solid interface. The crystal interface is surrounded by a boundary layer (as deboundary layer theory scribed by VOLMER’S [7.40]) in which molecules or atoms are weakly bonded to the crystal. The molecules remain mobile, and bond to the free lattice site which is most favorable in terms of energy. In the second step, according to VOLMER, mobile solid particles in the boundary layer are integrated into the crystal lattice [7.41, 7.421.

mean crystal growth rate (m/s) dimensioned constant (m/s) dimensionless constant relative velocity between the crystal and solution

The total growth resistance is therefore a summation of the diffusion and flow resistances, and an interfacial resistance. l h e crystal growth is usually expressed in practice by the increase of mass with time, dS/dT dS

-=

dt

c,

*

A , . ( X - X,)”

(7-59)

where C,,, is the crystal growth rate, A , is the total surface of the crystals and K describes the order of the “crystal growth reaction” (1 < H < 2). The crystal growth rate increases virtually linearly with the degree of supersaturation of the solution (Fig. 7-24). The rate increases with temperature and relative velocity, w, and decreases with increasing viscosity of the solution. Furthermore, the rate is influenced by the pH value and the impurities in the solution, and is different for each interfacial area with the same substances (see crystal form, crystal habit in [7.1, 7.31). The process of reducing supersaturation is described in detail in [7.38, 7.39, 7.44-7.471. To produce crystals with a mean diameter R and a narrow particle size distribution, control of the seed number is necessary. A crystal size density distribution, i. e., a balance for the number of crystals in each size

7.2 Crystallization from a Solution

range has to be formulated as well as mass and energy balances (for example, see [7.39, 7.501). The crystal size density distribution expresses the change in the number density dn/dx in the solution, in the size range dx with time. The changes can be due to growth, loss and/or generation of crystals, a change in the balance volume, and charging or discharging. If the nucleation rate and the crystal growth rate are known for a defined operating mode of the crystallizer, a crystal size distribution may be proposed from the crystal size distribution I7.391. The ratio of crystal growth rate rg = dx/dt, and the nucleation rate dn/dt, or B,, is plotted as the crystal size or grain index 5 against the supersaturation AX, as shown in Fig. 7-24 [7.52]. If the grain size of the product is unimportant, 5 may be small. The crystals may come from a solution of relatively high supersaturation, in which case mainly fine grains with large specific interfacial areas are produced. The crystallizer is, therefore, simple and small. If, however, coarse crystals with a small size distribution are required for handling and drying purposes, the solution supersaturation must be kept low. The required crystal interfacial area must therefore be large to reduce the supersaturation. The number of nuclei and the crystal growth must be controlled by controlling the supersaturation of the solution, which leads to an expensive crystallizer design. Important data for the design of technical crystallizers are found by experimental investigation of the crystallization kinetics. Different experimental setups [7.38, 7.491 on a laboratory or pilot scale are used under conditions as close as possible to the operation conditions, for example:

MSCPR crystallizer (mixed suspension, classified product removal) with homogeneous mixing and classified product discharging CSCPR crystallizer (classified suspension, classified product removal) with the crystal phase and the circulated mother liquor separated and classified product discharging - Moreover, scale up of the data obtained is virtually impossible, mainly due to different hydrodynamic conditions in the pilot-scale plant and the technical plant, impurities in the technical crystallizer which influence nucleation and crystal growth, and different heat transfer coefficients caused by different degrees of incrustation (or fouling) in pilot-scale and technical plants. Different models (plugflow, macromixing, etc.) are discussed in 17.381 which consider deviations from ideal behavior in an MSMPR crystallizer. Results of scale up for this type of crystallizer are given in [7.48].

7.2.4 Design of Crystallizers for Mass Crystallization from a Solution A variety of crystallizers are available for mass crystallisation from solutions on a technical scale. These offer different crystal yields and various crystal qualities (purity, grain size, grain size distribution), which depends on the design, the operating conditions and the crystallization process. Classification, based on the operating mode, gives: 0

Batch crystallizer MSMPR crystallizer (mixed suspension, mixed product removal) with homogeneous mixing and mixed product discharging

511

Batch crystallizers for small product outputs, equipped with stirrers or guiding tubes, and circulation pumps. The agitation equipment must be correctly operated in the sense of controlling the cooling or evaporating rate, and they must be

512

7 Solvent Evaporation, Crystallization

operated in the allowable supersaturated metastable range 0

Continuously operated crystallizers with lower limit range from approximately 80-400 kg/h, have the advantages of a high space-time yield, a lower degree of operating effort and a constant product quality over batch operation

Individual type classes have to be distinguished in continuous operations, depending on the range of crystal grain distribution. Increasing the product quality requirements further complicates the design of crystallizers. Most continuous crystallizers within these classes are suitable for most crystallization processes (evaporation crys-

tallization, cooling crystallization, vacuum crystallization and reaction crystallization). Table 7-5 gives a brief overview of the basic design types used in practice; the principle of classification is the form of contact between the solution and the crystals and the way that product is discharged. Frequently used crystallizer types are assigned to individual crystallization processes in Table 7-6. Design forms listed in Tables 7-5 and 7-6, are explained in more detail in References [7.1, 7.3, 7.4, 7.9, 7.511. Table 7-7 only gives a brief overview of different crystallizer designs with respect to quality of the produced crystals and the controlling criteria: supersaturation, nucleation and crystal phase. Supersaturation of

Table 7-5. Crystallizer types [7.51]’).

Magma Crystallizer without discharge of clear mother liquor (MSMPR)*)4, Simplest type of crystallizer with crystal suspension conveyed by means of mixer or pump, oversaturation is developed from the suspension. 0 Internal Forced Circulation Agitated crystallizer (single mixed vessel or horizontal crystallizer with multiple mixed cells) Draft tube crystallizer 0 External Forced Circulation (FC forced circulation) Crystallizer with external heat exchanger and circulation pump Magma Crystallizer with discharge of clear mother liquor (MSMPR)4) Crystal and solvent residence time are independently adjustable, separation of crystal slurry and suspension by sedimentation depending on the upflow velocity in the settling section. 0 Double propeller crystallizer (DP) 0 Draft-tube-baffle crystallizer (DTB) 0 Vortex crystallizer Classified-Suspension Crystallizer (CSCPR)3) 4, Suspension takes place in a liquid fluidized bed. Crystals are separated from circulated solution by sedimentation, thus the circulation pump only circulates crystal-free solution to avoid secondary nucleation and crystal attrition. Generation of coarse and uniform crystals. 0 Type “Krystal” (“Oslo”) Type “Messo” ’)

2, 3, 4,

Crystallizer schematic, see Table 7.7. MSMPR Mixed suspension, mixed product removal. CSCPR Classified suspension, classified product removal. The more complex the crystallizer design the more possibilities there are to control the crystal size distribution. Secondary nucleation rate increases with increasing complexity of the design.

7.2 Crystallization from a Solution

513

Table 7-6. Crystallizer types and corresponding crystallization processes. Solution Crystallizer 0 Cooling Crystallizer - Cooling crystallizer with evaporation cooling (direct solution cooling by evaporation of solvent in a gas flow). Tank crystallizer, pan crystallizer, shaking crystallizer (Wulff-Bock crystallizer), crystallizing rolls, tube crystallizer with and without internals, hose-type crystallizer. - Cooling crystallizer with cooled jacket (direct cooling of solution or suspension at the cooled crystallizer wall). Vessel crystallizer with cooled jacket, installed cooling elements or rotating cooling system for continuous and discontinuous operation, scraped cooler, rotary kiln crystallizer, disc crystallizer, double-walled jacket crystallizer (Voltator), drum crystallizer. 0 Evaporation Crystallizer Pan crystallizer with heat exchanger, tank crystallizer with heated jacket or internal heat exchangers and agitator, evaporator with natural or forced circulation agitator, single or multistage units, thin-film evaporator, evaporator with submerged burner. 0 Vacuum Crystallizer Agitated tank crystallizer for continuous and discontinuous operation, crystallizer with forced suspension circulation, horizontal crystallizer and multistage, crystallizer with agitator and controlled flow (DTB), vortex crystallizer according to Standard-MESSO. 0 Classifying Crystallizer Conical cooling crystallizer with settling zone according to Howard, vibrated crystallizer to prevent build-up of solid on the heating elements, fluidized-bed crystallizer system “Krystal” with cooling, evaporation or vacuum crystallization, DP-crystallizer system Escher-Wyss-Tsukishima, evaporation crystallizer with carrier gas by Robinson, vortex crystallizer Standard-MESSO. 0 Reaction crystallizer (combination crystallizer chemical reaction, for example, precipitation reaction). Spray crystallizer for melt and solution crystallization.

-

the solution is carried out by expansion cooling a n d flash evaporation under vacuum. With other supersaturation processes, such as pure cooling or solvent evaporation, the same criteria are also valid a n d so it is only necessary to add a n additional external or internal heat exchanger to the respective crystallizer design. The following requirements are for all crystallizers and are derived from the discussion in Chapter 7.2.3: 0

Large exchange areas between the crystal phase a n d the solution to intensify the reduction of supersaturation

0

0

0

0 0

A high crystal growth rate is achieved by taking advantage of the maximum possible supersaturation A sufficient relative velocity between the crystals and the solution to maintain the transport of solid from the solution to the crystal, and, therefore, to maintain the crystal growth Sufficient heating or cooling area is maintained by mechanical cleaning or by using a high solution velocity a n d thereby avoiding incrustation Low operation a n d investment costs Low space requirements

514

7 Solvent Evaporation, Crystallization

Table 7-7. Typical crystallizers using flash evaporation for solution supersaturation classified by supersaturation control, nucleation control and crystal bed control. Representation according to Mannesmann Engineering AG, Messo-Chemietechnik. Crystallizer type

Principles, characteristic data

Agitated crystallizer with uncontrolled solution supersaturation

Agitated crystallizer for continuous and discontinuous operation, feed temperature d, , cooled to 19,, 19, is also exit temperature of crystal slurry, nucleation and crystal growth at randomly adjusted equilibrium Crystallizer performance is limited by solution boil over. Solution throughput: 0.1-3 m3/h Crystal size: 0.1-0.5 mm Low crystallizer performance and large vessel volume increase the crystal size 1 Body, 2 Agitator, 3 Crystal slurry overflow

1x1

Vacuum

Crystal' slurry

Forced-circulation crystallizer with controlled solution supersaturation

7tle.r

solution

Horizontal vacuum crystallizer with controlled solution supersaturation and limited crystal phase control

Vacuum

Clear solution

1

4

Crystal

To control solution supersaturation feed is introduced to the solution circulated by a circulation pump. Crystal slurry is mixed with feed (fresh solution). The mixing temperature d2 has to be adjusted by the amount circulated so that supersaturation of the solution at the boiling surface does not cross the border of the metastable area. Supersaturation of the circulated solution at the boiling solution level occurs through cooling. Crystal size: 0.3-0.8 mm Disadvantage: Imperfect control of nucleation 1 Body, 2 Circulation pump, 3 Crystal slurry overflow Total supersaturation by division into 7 individual stages (11 stages possible) with 7 increasing small temperature steps in 7 small concentration steps. Compared to one-stage operation supersaturation in each stage is formed by smaller circulation amounts with less attrition and less nucleation. Larger residence of the crystals than the solution at appropiate design of 2 the crystal slurry jack legs. Solvent throughput: up to ~ 5 0 m3/h 0 Crystals produced: up to 50 t/h Crystal size: 0.3-0.8 mm 1 Body, 2 Circulation pump, slurry 3 Crystal slurry overflow, 4 Jack leg

7.2 Crystallization from a Solution

515

Table 7-7.(continued) Crystallizer type

Principles, characteristic data

Forced-circulation crystallizer with controlled solution supersaturation and controlled crystal bed (Oslo crystallizer)

Temperature increase caused by the feed is limited by the controlled circulation of clear mother liquor. Thus supersaturation of the circulated solution is kept slightly in the metastable zone during cooling, no new nuclei are generated. Supersaturated solution flows directly to the bottom of the crystal bed. Supersaturation is transferred to crystals (crystal growth) during upflow of solution through the crystal bed. Due to a controlled discharge of clear mother liquor the crystal slurry is concentrated up to the flowability limit. Solution throughput: = 500- 1000 m3/h by connecting several stages in series Crystals produced: up to 100 t/h Crystal size: 0.6-3 mm 1 Crystallizer, 2 Circulating pump, 3 Crystal slurry jack leg, 4 Clear mother liquor valve

Forced-circulation crystallizer with control The crystal growth zone operates according to the of solution supersaturation, crystal phase principle of the Oslo crystallizer. Circulated solution circulated via suction nozzle 6 is mixed with superand nucleation (Messo crystallizer saturated solution from the top cooling zone passing down through the inner tube 2 and the outer tube 7. In the bottom section of the crystallizer this mixture is guided first to the outside and to the crystal growth zone 11, here supersaturation is transferred to the growing crystals. By means of control valve 13 clear mother liquor from settling zone 10 is discharged via exit fitting 9 a suspension with controlled crystal concentration is produced, 8. Fresh solution or feed is introduced at the bottom 1, to the cycle. Crystal slurry is classified (crystal size classification) by the circulation amount of the outer cycle which is controlled by the suction nozzles. Small crystals from solutlon I the solution are entrained from the crystal bed and 1 Feed inlet fitting recycled to the inner circulation driven by the pro2 Draft tube peller pump 3, where supersaturation is reduced due 3 Propeller pump to degradation of crystal fines. Different crystal con4 Lower skirt baffle centrations and sizes in inner and outer suspension 5 Driving nozzle cycles are caused by different circulation amounts in 6 Suction nozzle each cycle. With known crystallization behavior the 7 Upper skirt baffle cycle may be controlled in a way, that in the inner 8 Crystal slurry overflow cycle at a low suspension concentration and a rela9 Clear solution exit fitting tively high supersaturation nucleation occurs. Other10 Settling zone wise, in the outer cycle at a longer residence time 11 Crystal growth zone and smaller supersaturation crystal growth is favored. 12 Vapor fitting Crystal sizes: 0.8-4 mm 13 Control valve No crystal fines below 0.3 mm

516

7 Solvent Evaporation, Crystallization

If uniform coarse crystals are to be produced, an additional requirement is: 0

The control of the supersaturation with controlled nucleation and control of the crystal phase by using a classifying operation mode

a cooler or evaporator (Fig. 7-28), a mass balance gives

(F, + F,,,) . AXo=

s

(7-60)

where AXois the required reduction in the supersaturation to produce the crystalline product. AX, is the supersaturation related These requirements have to be consid- to F, + pm,given in kg solid/kg mother liered during the selection and design of quor. It is assumed that the total supersatucrystallizers. ration of the mixture F, + F,,, in the crystal bed, is converted into the crystalline product, If this is not the case, the left side of Eq. (7-60) has to be multiplied by a factor 7.2.5 Criteria for the Selection which is < 1. and Design of Crystallizers (The allowable supersaturation AXo must be experimentally determined (see Selection of a crystallization process, the Chapter 7.2.3). For a given crystalline prodtype and design of a technical crystallizer, a uct yield the required circulated suspenheat exchanger and other associated equip- sion rate Fmis given by Eq. (7-60), and is ment including circulation, equipment, etc., fixed by the supersaturation, which must be should ensure trouble-free long-term opera- smaller than the metastable range of supertion, limited incrustation of the inside ap- saturation. An estimated value for the alparatus walls, optimized economic produc- lowable supersaturation is given as AXo A tion of crystals at a required production Ac = 1 g/L. When determining FTn,it must rate and the required quality (purity, grain also be considered that the state point of size, grain size distribution). Figure 7-25 the mixture F, + Fm lies in the range of the shows schematically the course of action metastable solution). The supersaturation of the solution is necessary for crystallizer selection and dereduced as it flows through the crystal bed sign. Additional considerations for the selec- of the crystallizer. A sufficient crystal bed tion of a crystallizer are given in Fig. 7-26. height and solution residence time in the Figure 7-27 shows the specific crystal prod- crystal bed are required for this. Therefore, the required cross-sectional uct rate of CSCPR crystallizers, with fluidbed suspension. Modeling of crystallizers is area AQ of a cylindrical crystallizer is described in [7-661. The determination of the main dimensions of classifying crystallizers is briefly discussed in the following. For a comprehensive design method, see References [7.1, 7.39, 7.511. In a classifying crystallizer, such where em is the density of the supersatuas the Oslo and Messo crystallizer, a separa- rated solution and w is the superficial veloction effect is caused by the different sedi- ity related to the free cross-sectional area mentation behavior of different grain sizes AQ. Therefore, the crystallizer diameter increases with decreasing supersaturation. in the upflowing solution. The specific crystalline rate (output per If fresh solution F , is mixed with circulated mother liquor Fmand supersatured in unit cross-sectional area) s/AQof classify-

s.

s

7.2 Crystallization from a Solution Feed properties (amount, composition, thermal behavior, corrosive properties etc.)

Production capacity (produced

I

t

Operating mode and type of peripheral equipment

Physical and chemical feed data (temperature, density, viscosity, saturation temperature, dissolved substance fraction etc.)

Physical and chemical data of mother liquor (especially solubility function, metastable oversaturation zone, boiling point increase, density, viscosity)

crystals (nucleation and growing rate, crystallization enthalpy, density, terminal velocity etc.)

quality (purity, particle size, uniformness)

Product market price (+allowable investment and operating cost)

I

+

4Determination of crystallization process I -

517

4Operating mode selection -k

-

c Crystallizer type selection

-b

Balancing: Mass balance Heat balance Energy consumption Particle number balance

Design of crystallizer periphery

Fig. 7-25. Course of action necessary for crystallizer selection and design (schematic).

518

7 Solvent Evaporation, Crystallization

X X

X

X

X

x x x x

Circulation E Circulation I E Circulation 0 Draft-tube ? Draft-tube 7 Draft-tube

X

x x x x

x

X

2

-

x x x x

x x x

X

X

x x

X X

x x

X

x x

crystallizer crystallizer crystallizer crystallizer crystallizer with FCD crystallizer with FCD

m E d Draft-tube crystallizer with FCD .J I

Z

E

Oslo crystallizer O S I ~crystallizer

0s10 crystallizer

Evaporation Vacuum cooling Surface cooling Vacuum cooling Evaporation Vacuum cooling Surfoce cooling Evaporation Vacuum cooling Surface cooling

Fig. 7-26. Selection of crystallizer type. Representation according to WOHLK,Mannesmann-Anlagenbau AC, Messo-Chemietechnik. MSMPR Mixed suspension, mixed product removal MSCPR Mixed suspension, classified product removal CSCPR Classified suspension, classified product removal FCD Fine crystal discharge

-r -E

500

-

LOO

2Yl 0)

I

I

I

I -A c = 3 g /I

c

2 300

5

F

150

%

a

c 0

.-u

In n

I

I

-Ac = 2 g/l

d c :1 g/l

100

0

10

20

30

Upward flow velocity [mm/sl-

LO

Fig. 7-27. Specific crystal product rate of CSCPR crystallizers as a function of the upward flow velocity and allowable oversaturation Ac. Representation according to WOHLK and HOFMANN [7.51].

7.2 Crystallization from a Solution Mother

+

liquor

-

Fm r

"Fl CP

Fig. 7-28. Crystallizer cross-sectional area AQ and crystal bed height 2, of a classifying crystallizer. CR Crystallizer CB Crystal bed CP Circulating pump CV Cooler or evaporator for solution supersaturation

ing crystallizers, is controlled by the terminal velocity of the smallest crystal to be separated in the settling section, and therefore by the upflow velocity of the circulated solution. The larger the terminal velocity of the particle cut size and the larger the allowable supersaturation, the larger is the value of %!AQ.For example, for production of 10 ton/h potassium chloride in an Oslo crystallizer, with a specific production rate s/Ag = 190 kg/(m2h) a cross-sectional area AQof about 53 m2 is required, which corresponds to a diameter of 8.2 m (see also Fig. 7-27). For a Magma crystallizer with removal of the clear solution and S/AQ7 500 kg/(m2h), a diameter of only ca. 5 m is required for the same production rate. However, the crystal size in an Oslo crystallizer is approximately 2 mm, whereas in the Magma crystallizer it is ca. 1 mm [7.51].

519

The mass mk of the crystals in the crystal bed is given in kg suspended solid/m3 bed volume, along with the bed density eb,

The larger the initial supersaturation AX,, the velocity w and the mean residence time t,, the larger the crystal bed height. Conversely, the height decreases with increasing bed density. For example, a higher velocity leads to a looser crystal bed and, therefore, to a lower bed density. With smaller grain sizes, the flow velocity must be lower than with coarse particles to ensure a sufficient bed density. AX, may now be assigned a larger value, since the supersaturation reduces more quickly due to the larger specific surface area of the fine grain. For further discussion of the variables on the right side of Eq. (7-63), which are mainly determined experimentally, see [7.1]. Estimated data for different applications are also given there.

7.2.6 Freezing In freezing processes the solvent is crystallized by cooling the solution. The remaining concentrated solution of the dissolved substance is usually obtained after multistage separation of the crystal phase by squeezing, centrifugation or removal of floating crystals.

520

7 Solvent Evaporation, Crystallization

Due to the low melting point of the solvent and the course of the liquidus curve (“ice curve”) in the phase diagram (Figs. 1-39 and 7-29), freezing processes are usually operated at low temperature. This enables a gentle concentration of the solution. Freezing is possible in all solvent-solute systems forming an eutectic, but as yet, is mainly only applied to aqueous solutions. A flat path of the liquidus line, from the melting point of the solvent to the eutectic point, increases the yield of solvent crystals for a given cooling temperature range. The larger the distance between the solvent melting point and the eutectic point, the better the concentration of the solution by freezing. If solution flow F , , given in kg/h of initial mass fraction w, of dissolved solid, is cooled from the initial temperature LS, (Fig. 7-29) to the temperature r9, on the liquidus line, the solvent starts to crystallize. The solution becomes enriched in the dissolved substance, whilst the equilibrium

temperature drops according to the route of the liquidus line. If the freezing process ends at temperature c9,, the solid mass fraction of the solution is increased to w2.According to the lever rule for the phases, the ratio of the fraction of solution crystals (“ice”) to the remaining quantity of solution, is the same as the ratio of the distances BC to AB, i. e., (7-64) With the remaining solution flow as

F2 = F, - sL the solvent crystal yield process is given by

s, = F 1 (I.

(7-65)

s, for the freezing

-2)

(7-66)

The phase equilibrium expressed by the liquidus line, is usually not reached in practice, since the frozen solvent contains some solute. The solvent crystal yield is thus not exactly the value calculated from Eq. (7-64) or Eq. (7-66), [7.1]. The heat transfer during freezing is directly carried out using heat exchangers or by evaporation of solvent under a vacuum [7.53]. Inmiscible auxiliary substances with a low boiling point are sometimes added to the solution, acting as a heat transfer agent in vacuum evaporation.

sL

7.2.7 Fractional Crystallization from a Solution Fig. 7-29. Phase diagram: solvent freezing.

LC Liquidus curve SC Solubility curve SOC Solidus curve ~5

w

Temperature Mass fraction of dissolved substance

In fractional c r y ~ ~ a ~ l i z ~processes, tion a crystal phase containing two or more components is brought into intensive contact with a suitable mother liquor several times,

7.3 Crystallization from a Melt

and is recrystallized at the same time. Fractionation begins with the contact of pure solvent with the crystal phase and is usually a discontinuous operation due to specific characteristics of the mother liquor and the crystal mixture [7.1, 7.31. Continuous fractionation is carried out in countercurrent flow. For the continuous fractionation of a solution containing two dissolved substances, a split column is used (Fig. 7-30). In the upper column section, the crystallization column, the less soluble component crystallizes during the countercurrent contact of the mother liquor and the crystal. In the lower section, the concentrating or enrichment column, the more soluble component is favorably extracted from

’c-

UR

Esr Bottom product ‘crystals

Fig. 7-30. Fractional crystallization from solution in countercurrent columns. CC Crystallization column ES Enrichment section EC Evaporator crystallizer FI Filter UR Upper reflux LR Lower reflux

521

the crystalline phase, in which the less soluble component is enriched. The enriched crystal phase is then discharged from the bottom of the column. Some of the bottom product crystals are dissolved and fed to the enrichment column as lower reflux. The solution withdrawn from the top of the crystallization column is enriched with the more soluble component. After evaporation of solvent, some of the crystals flow back to the column as upper reflux and act as seeds to initiate crystallization of the less soluble component. The required number of stages for fractional crystallization are determined by methods given in [7.1]. The PonchonSavarit and the McCabe-Thiele diagram are both commonly used. Conversely, in recrystallisation processes, only the crystalline phase is purified by further treatment, not the mother liquor.

7.3 Crystallization from a Melt Crystallization from a pure melt consisting of only one component is a simple formulation process for the preparation of the sales product, and may be carried out on watercooled steel belts or on internally cooled drums which dip into the melt (Fig. 7-31). The operating conditions for the crystallization unit are given by the path of the melting point pressure curve of the component to be treated (see Chapter 1.4.5.2). If the melt is a mixture of two or more components with different melting points, the crystallization is then a true partial thermal separation process of the mixture, once the crystalline phase and remaining melt have been separated. The degree of separation of a molten mixture by partial crystallization depends on the equilibrium behavior (see Chapter 1.4.5.4). If the mixture components are

522

7 Solvent Evaporation, Crystallization

Liquid feed

\iI

TF

Crystals

CP

a)

Liquid feed

I

1

Fig. 7-31. Belt crystallizer (a), drum crystallizer (b). Representation according to Sandvik Conveyor GmbH. Schmiden. CD Charging device, liquid feed SB Steel belt, cooled eutectical composition (mixture type 11, CA Cooler, cooling agent Fig. 1-43). CP Circulating pump, cooling agent In comparison to distillation, melt crysCR Crystals receiver tallization at at low temperature is a therCT Cooling agent trough DR Cooling drum cooled intern mally gentle separation process and favors MT Melt trough separation of mixtures which contain comSR Scraper ponents in a narrow boiling point range or TF Thin film layer of freezing melt azeotropic mixtures, and for purification of

completely miscible in the liquid and the solid state, virtually complete mixture separation is theoretically possible (mixture type I, Fig. 1-43). Therefore, the separation apparatus must have a minimum number of separation stages. If the mixture components are only miscible in the liquid state, and not in the solid state, then partial crystallization leads to a virtually pure component product and a remaining mixture of

thermally sensitive substances [7.54,7.651. Due to the low operating temperature, and the melting enthalpies rather than the evaporation enthalpies of the mixture components, crystallization from a melt requires a much smaller energy input. The process which is used on a technical scale for the crystallization from a melt, to separate and purify mixtures, may be described as follows: 0

With discontinuous operation only simple equipment designs are required. Crys-

7.3 Crystallization from a Melt

0

tal layers continuously grow at the walls giving simple separation from the remaining melt (tube bundle crystallizers, falling film crystallizers, bubble column crystallizers, cascades of crystallizer drums) With mainly continuous operation, the entire melt is converted to a crystal suspension in the crystallizer by cooling, and is followed by a mechanical separation of the solid and remaining melt (mixed vessel crystallizer, cooling disc crystallizer, forced circulation cooler with scraper, crystallizer columns)

The second process group enables multistage mass transfer in countercurrent flow of melt and crystal phase and therefore a good separation efficiency. The main disadvantage is that the capital expenditure for equipment and conveying devices to control the relative flow of the crystalline phase and melt phases is higher. With the first group, multistage crystallization is only possible in a cascade or two-cycle mode by repeating the crystallization and melting steps (see Fig. 7-6). Figure 7-32 shows three basic designs of common continuous countercurrent column crystallizers from a melt. A mixture of crystals and melt, previously produced by partial crystallization, is charged to a “Philips” column (Fig. 7-32a). Some of the remaining melt, enriched with the lower melting component, and the upflowing reflux from the fractionating section are discharged under pressure from the column through filter slits. The crystals sediment from the filtration section to the fractionating section and are then enriched with high-melting product in countercurrent flow with upflowing melt. After reaching the melting zone, the enriched crystals melt. Some of the melt leaves the column as a high-melting product of high purity. The remaining melt is the melt reflux to the column, which is pulsated in the fractionating

523

zone to obtain more intensive mass transfer. The “Phillips” pressure column is particularly suitable for the treatment of systems forming an eutectic point. In the countercurrent column of SCHILDKNECHT (Fig. 7-32b), the melt is fed centrally to the column. Mass and heat transfer take place between the crystal phase which is transported upward by a rotating conveying spiral, and the downflowing melt. The lower melting product is discharged at the bottom of the column and the higher melting product at the top of the column. Figure 7-32c shows a column with centrally fed melt and with upper and lower reflux. Such a column may be operated adiabatically and can be equipped with different internals. The method to determine the required number of stages for countercurrent crystallization columns is described in [0.1] and [7.1]. Table 7-8 gives a brief overview of melt crystallizers with references to further literature. Process examples to purify substances and to separate mixtures by crystallization from a melt, are given in [7.54]. If separation processes operating at their highest separation efficiency are combined, energy is usually saved. For example, for an azeotropic system which forms mixed crystals in the solid state, separation is performed by combining rectification and melt crystallization above the azeotropic point (Fig. 7-39). The resulting mixtures of the rectification D, and D, are then treated as mixture G by melt crystallization. A crystalline phase K and a remaining melt R are obtained. The melt K is then separated into virtually pure components A and B by rectification and into the mixtures D, and D, after appropriate preparation or by adding mixture F. (Additional examples of rectification combined with melt crystallization or extraction with melt crystallization, are given in [7.54]).

524 Liquid fetd

-i

7 Solvent Evaporation, Crystallization a1

b)

MS

High melting product

-

2-

I

1

l

I

melting product

-cc

Liquid feed

r

sc

product

l l Low melting product

Fig. 7-32. Different types of countercurrent column crystallizers. a) End-fed pulse column, “Phillips” type b) Countercurrent centre-fed column with spiral conveyor, Schildnecht type c) Countercurrent centre-fed column with upper and lower reflux FR Freezer FI Filter CC Countercurrent crystallization column FS Fractionation section product is then a single or multicomponent MS Melting section PU Pulse unit crystal. SC Spiral conveyor Desublimation is a pure molding process ME Melter if the vapor phase contains only one comCS Crystal slurry ponent. The operating pressure and temperUR Upper liquid reflux ature of the desublimation unit are fixed by LR Lower slurry reflux

7.4 Crystallization from a Vapor Phase, Sublimation and Desublimation Substances are directly converted from the solid phase into the vapor phase in sublimation processes. Conversely, substances are “condensed” directly to the solid phase from vapor in desublimation processes. The

the sublimation pressure curve of the component to be treated (see Chapter 1.4.4.1). At least the heat of sublimation has to be removed during sublimation. If the vapor contains several components, a partial desublimation leads after separation of the solid phase from the vapor phase to a partial separation of the original vapor mixture. With simple sublimation for substances which sublime easily at the operating temperature (substances with high sublimation pressure), a heated sublimer and a cooled desublimer are sufficient for sublimation/ desublimation cycles.

7.4 Crystallization from a Vapor Phase, Sublimation and Desublimation

The component to be purified sublimes from the contaminated feed by supplying heat to the sublimer and is then separated by removing heat in the following desublimer, giving purified product (Fig. 7-40a). If the vapor pressure of the sublimate is not high enough for simple sublimation, the process must be operated under vacuum or with an auxiliary component used as a

525

carrier gas (entrainer). With vucuum sublimation (Fig. 7-40b), the total pressure in the sublimer must be lower than the sublimation vapor pressure of the treated component, at the operating temperature. With entrainer sublimation (Fig. 7-40c), it is sufficient to reduce the partial pressure of the component to be purified below the sublimation pressure at the operating tempera-

Table 7-8. Selected crystallizer designs for crystallization from a melt. Representation according to RITTNERand STEINER[7.54]. Usually intermittently operated crystallizer with crystal layer formation in the cooling area 0

Tube bundle crystallizer, design Proabd [7.56]

Heater

:i

Fig. 7-33. Tube bundle apparatus, usually tubes with fins. During crystal layer formation cooling agent flows through tubes, whereas during exudating of impurities from the crystal layer heating agent flows through the pipes. Impurities may be discharged sectionwise or while melting the crystals.

i-

Cooling section

~~

\-

\

Collector--tube

-

r (melt)

Residuol melt/ s w e a t s oil/ pure product

(continued next page)

526

7 Solvent Evaporation, Crystallization

Table 7-8. (continued) Falling-film crystallizer by Sulzer-MWB i7.571

0

Fig. 7-34. Long-tube apparatus (tube diameter 50-75 mm, tube length 12 m) with falling-film flow on both cooling agent and product side, temperature control of cooling agent thus growth of crystals on the tube wall, after several circulations of the melt and ending of the crystallization period, crystals are molten and discharged (see also Fig. 7-6).

Cooling or heating agent

Residual melt/ sweats oil/ pure product

0

Crystallizing rolls cascade [7.64]

Squeeze roll Crystallizing rolls

Feed lmelt)

Scraper Melt pan

' Residual meH

Pure product

Fig. 7-35. Crystallizing rolls cascade combined with squeeze rolls if necessary, the higher melting component is preferentially deposited at the surface of the crystallizer rolls. Higher separation efficiency (yield) and purity are obtained by melting the purified crystals in the next higher purity stage and passing the residual melt in the stage of lower purity to become crystallized again. From the stage of highest purity one part of the final product is used as a melt reflux (countercurrent flow of crystals and melt).

7.4 Crystallization from a Vapor Phase, Sublimation and Desublimation

527

Table 7-8. (continued)

Crystallizer for mainly continuous operation with formation of a crystal suspension 0

Agitated crystallizer Fig. 7-36. Formation of a crystal suspension in an agitated vessel by jacket cooling or by direct contact of melt and an inert cooling agent, i. e., Cot.

F e e d lmeltl

o n

Cooling agent 1i.e.. C0,I

3 0

Jacket cooling

0 b

' 0

0 0

Crystal suspension

Direct coollng (cooling agent)

Scraped chiller [7.59] Fig. 7-37. Horizontal tube-like crystallizer with jacket cooling. Inside the tubes flexible wiper systems are arranged to remove crystal depositions, control of crystallization conditions by temperature and circulated amount.

Crystal suspension

Feed (melt1

Cooling agent

Chilled-disc crystallizer [7.60] Fig. 7-38. Jacket-cooled crystallizer with cooling disc mounted on a rotor to cool stepwise and to considerably avoid backmixing of melt or crystal suspension.

Cooling agent -eed lmeltl

GCrystal suspension

F Cooling agent

0

Crystallization column Countercurrent flow of melt and crystals (Fig. 7-32). Pressurized columns with crystal or melt transport by external pumps or ram, columns with mechanical forced conveyors, i. e., mixer, screw, spiral, etc.

528

a'

7 Solvent Evaporation, Crystallization

1

L

a)

SE

r---+----I

HA=*:

Waste gas 4

\&IfCA Residue

e:i:", feed

x u r e solid product

CA

b)

cn

hl

Residue

crude feed

$ P u r e solid product

Fig. 7-39. Combination of distillation and melt crystallization to separate a binary system with mixed crystal formation and azeotropic boiling behavior. Representation according to RITTNER,STEINER [7.541. a) y, x-Diagram: McCabe-Thiele diagram b) 7; x-Diagram: Solid-liquid phase diagram

ture; the partial pressure difference with the operating pressure is composed of the partial pressure of the auxiliary component used as the entrainer. In sublimation and desublimation processes mainly direct heated or cooled chambers or drums, with scrapers and conveying devices and fixed or rotating internals are used (Table 7-9). Partial sublimation of a gas mixture is frequently carried out in chambers equipped with finned-tube heat exchangers. During the desublimation period, the cooling agent flows through the heat exchanger, whilst in the following melting period, the heat exchanger is charged with a heating agent. Therefore, a continuous desublimation process for a gas phase requires at least two chambers.

Fig. 7-40. Different types of sublimation processes. a) Simple sublimation b) Vacuum sublimation c) Entrainer sublimation SU Sublimer GC Gas cooler BL Blower CA Cooling agent SB Sublimate DS Desublimer VP Vacuum pump H E Heater H A Heating agent ES Entrainer + solids

7.4 Crystallization from a Vapor Phase, Sublimation and Desublimation

529

Table 7-9. Selected sublimer designs. Representation according to MATZ [7.2]. Tray sublimator [7.62] Fig. 7-41. Vacuum sublimator with fixed trays heated by steam or other heating agents and cooled desublimation areas. Sublimed vapor crystallizes at the desublimation area, where solid product is removed by means of rotating brushes. a Cooling agent exit; b Cover and body;

c Sublimation tray; d Desublimation area; e Rotating brushes; f Vacuumproofed rotary seal; g Cooling agent inlet; h Coupling; i Drive.

0

Turbine sublimator [7.63] Fig. 7-42. Entrainer sublimator with rotating disc on top of each other an1 divided into sections to receive solids feed. Entrainer or carrier gas is recycled by a turbine. Entrainer flows countercurrently to the solid which is passed from tray to tray from the top to bottom. Sublimed vapor is picked up by the entrainer. Separation of entrainer and vapor takes place in the desublimator. The entrainer then recycles to the sublimator. Diameter of sublimer: 1.8-10 m Height: 1.8-20 m. a Feed; b Rotary-vane feeder; c Hot filter; d Fan; e Desublimer; f Heater; g Wiper; h Levelling; i Turbine sublimer;

k Heated base; 1 Collector for residue; m Heater for recycled air.

530

7 Solvent Evaporation, Crystallization

Easier

~.-scsublimating product

u!l

GE

Heavier

I sublimating product

L--{

’s

H A ;> J

‘i

Fig. 7-43. Fractional sublimation in a countercurrent column with solid and gaseous entrainer. DS Desublimer SC Sublimation column SU Sublimer GE Inert gaseous entrainer SE Inert solid entrainer HA Heating agent CA Cooling agent

The separation of a mixture containing sublimable components is carried out by fractional sublimation in countercurrent columns. The finely distributed solid phase trickles down the column countercurrently to the upflowing sublimate vapor (Fig. 7-43). For the treatment of binary mixtures, the lighter sublimable component is withdrawn from the top and the heavier sublimable component is withdrawn from the bottom of the column. Some of the lighter sublimable component is used as solid reflux at the top of the column L7.551. Sublimation/desublimation processes are rarely used due to the high costs and the difficulty and complexity of the operation of handling the solid phase in the process units. However, it is’a thermally gentle purification process for the separation of desublimable components from a gas mixture and sometimes a separation of mixtures composed of different sublimable components.

References

References [7.I ] MATZ, G.: Kristallisation. Grundlagen und Technik. Springer, Berlin 1969. [7.2] MATZ,G.: Chem. Ing. Tech. 50 (1978) 1, 13-23. [7.3] MULLIN,J. W. : Crystallisation. Butterworth, London 1972. [7.4] VAN HOOK, A. : Crysta/lisation. Theory and Practice, Reinhold Publishing, New York 1961. [7.5] NYVLT, J. : Industrial Crystallisation from Solutions. Butterworth, London 1971. (7.61 NYVLT,J., et al.: The Kinetics of Industrial Crystallization. Elsevier Science Publishers, Amsterdam 1984. S. J., and G R O ~ S C H O L TP.E A. N, [7.7] JANCIC, M. : Industrial Crystallization. Reidel Publishing Comp., Dordrecht 1984. [7.8] ZIEF,M., and WILCOX,W. R.: Fractional Solidification. 2 Vols. Dekker Inc., New York 1967. [7.9] BAMFORTH, A. W.: Industrial Crystallisation. Hill, London 1965. [7.10] MULLIN,J. W., et a].: Industrial Crystallization. Plenum Press, New York 1976. [7.11] MATZ,G. : Fortschr. Verfahrenstech. 22 (1984) Sec. C, 305-323. [7.12] NYVLT, J. : lnd~strial Crystalfizution. Verlag Chemie, Weinheim 1982. [7.13] MATZ,C.: Chem. fng. Tech. 55 (1983) 1, 72-75. [7.14] VENER,R. E., and WOMPSON, A. R.: Ind. Eng. Chem. 42 (1950) 464-467. H. M.: New [7.15] FINDLAY, R. S., and SCHOEN, Chemical Engineering Separation Techniques, A bschnitt Adductive crystallisation. Interscience Publishers, New York 1962. [7.16] MESSING,T.: Chem. Ing. Tech. 42 (1970) 18, 1141- 1148. 17.171 MATZ, G.: Chem. Tech. (Heidelberg) 5 (1976) 6, 201-206. [7.18] MATZ, G., and KAUFHOLD, G.: Chem. Tech. (Heidelberg) 8 (1979) 8, 373-376. [7.19] ADAMSKI, T.: CAV 3 (1971) 62-70. [7.20] BILLET,R. : Verdampfertechnik, Bibliographisches Institut, Mannheim 1965.

531

[7.21] BILLET,R. : V2rdampfungund ihre technischen Anwendungen, Verlag Chernie, Weinheirn 1981. [7.22] WAGNER,J.: Chem. Ind. 35 (1983) 4, 212-216. C. : Fortschrittsber. VDI [7.23] MOSTOFIZADEH, Z., Reihe 6, 70 (1980). [7.24] GREGORIG,R. : Warmeaustausch und Warmeaustauscher Sauerlander, Aarau 1973. [7.25] KERN, D. Q.: Process Heat Dansfer. McGraw-Hill Kogakhsha Ltd., Tokyo 1950. [7.26] RANT, 2 . : Verdampfer in Theorie und Praxis. Steinkopff, Dresden 1977. H., and SLIPCEVIC, B.: Chem. [7.27] SCHNELL, Ing. Tech. 56 (1984) 6, 441-446. [7.28] BOHNET,M.: Chem. Ing. Tech. 57 (1985) 1, 24-36. [7.29] WOHLK,W., and LINSTAEDT, W. : BWK 35 (1983) 6, 691. [7.30] SCHREINER,H. : Verfahrenstechnik 3 (1969) 7, 301-308. [7.31] HOFFER,K.: Dissertation, T H Karlruhe 1962. [7.32] HOFFER,K.: Chem. Ing. Tech. 35 (1963) 147- 154. [7.33] VDI-Warmeatlas. VDI-Verlag Diisseldorf, 1974. [7.34] SCHLUNDER,E. U. : Vt-Hochschulkurs Thermische ~ r f a h r e n s t e c h ~ ~Krausk, kopf-Verlag Mainz 1972. (7.351 NAGEL,0.: Dissertation, T H Karlsruhe 1963. [7.36] WOHLK,W., and HOFMANN, G.: Chem. Ing. Tech. 52 (1980) 11, 898-900. A.: Chem. Ing. Tech. 49 [7.37] MERSMANN, (1977) 9, 679-691. [7.38] MERSMANN, A., BEER, W. F., and SEIFERT,D.: Chem. Ing. Tech. 50 (1978) 2, 65-76. [7.39] MERSMANN, A., and KIND, M.: Chem. Ing. Tech. 57 (1985) 3, 190-200. [7.40] VOLMER, M.: Z. Phys. Chem. 102 (1922), 267 -275. [7.41] KOSSEL, W.: Ann. Phys. 21 (1934), 457. [7.42] STANSKI,I. N.: Z. Phys. Chem. 136 (1928), 259. [7.43] MCCABE,W. L., and STEVENS, R. P.: Chem. Eng. Prog. 47 (1951), 168.

532

7 Solvent Evaporation, Crystallization

[7.59] SOEKADAR, K. : Verfahrenstechnik7 (1973) [7.44] MATZ,G. : Verfahrenstechnik 3 (1969) 5, 209. 191-198, and 7 (1973) 3, 76-81. [7.45] MATZ,G.: Chem. Ing. Tech. 45 (1973) 3, [7.60] Company report: Fa. Goudsche Maschinefabriek, Gouda. 101- 105. [7.46] MATZ,G.: C Z Chem. Tech. 3 (1974) 12, [7.61] SAITOH,S., IMURA,T., and KODAMA,K.: Dechema-Monogr. 66 (1971) 321. 429 - 434. [7.62] Company report: Fa. Leybold-HochvakuA. D.: [7.47] MOYERS,C. G., and RANDOLPH, urn-Anlagen GrnbH, Koln. A I C h E . J. 19 (1973) 6, 1098-1104. [7.48] Pmss, R., T ~ N G L E RT., , and MERSMANN, [7.63] NORD, M.: Food Manufact. 27 (1952) 452-457. A.: Chem. Ing. Tech. 57 (1985) 6, [7.64] CHATY,I. C., O’HERN,H. A.: A I C h E J. 536-537. 10 (1964) 74. [7.49] WOHLK,W.: Fortschr. Ber. V D I Z. Reihe [7.65] HEIN, K., and BUHRIG,E.: Kristallisa3. 71 (1982) tion aus Schmelzen. VEB Deutscher [7.50] DE JONG, E. J.: Chem. Ing. Tech. 54 Verlag fur Grundstoffindustrie, Leipzig (1982) 3, 193-202. 1983. [7.51] WOHLK,W., and HOFMANN, G . : Chem. Ing. Tech. 57 (1985) 4, 318-327. [7.66] MERSMANN,A., and KIND, M.: Rer. Bunsenges. Phys. Chem. 90 (1986) [7.52] MESSING,T.: Chem. Ing. Tech. 40 (1968) 955 -963. 16, 793-799. G. : CZ Chem. Tech. 3 (1974) [7.67] COULSON,J. M., RICHARDSON,J. F., [7.53] STEINBACH, BACKHURST, J. R., and HARKER,J. H . : 10, 363-368. [7.54] RITTNER,S., and STEINER,R.: Chem. Crystallization in Chemical Engineering. Vol. 2. Pergamon Press, Elrnsford, Oxrng. Tech. 57 (1985) 2, 91-102. [7.55] PREGER,M. : Aunereitungslech., 9 (1970) ford 1991. 551 -558. [7.68] MULLIN, J. W. : “Crystallization and Precipitation. Sublimation.” Ullmann’s [7.56] FR-PS 8 108 509 (1981), BEFS Engineering, Mulhouse Encyclopedia of Industrial Chemistrj [7.57] MAYER,M.: Verfahrenstechnik 8 (1974) Vol. B2, B3. VCH Verlagsgesellschaft, Weinheirn 1988. 221. K.: Chem. Ing. Tech. 55 [7.58] STOLZENBERG, (1983) 1, 45-46.

Thermal Separation Processes: Principles and Design Klaus Sattler, Hans Jacob Feindt copyright 0VCH

Verlagsgesellschaft m h H , 1995

8 Documentation and Calculation of Physical Characteristics

For the design engineer, it is usually very difficult to find in the range of interest the required physical characteristics of substances in a reliable and consistent form. These characteristics include: pressure, temperature and concentration dependence for the design of separation units. However, good data collections are given in 18.1 -8.121. These data collections complement particular literature of physical characteristics, cited in detail in Chapter 1. Additional collections of physical characteristics, are given in [8.13]. Reference 18.141 is a register of physical characteristics collected worldwide and compiled with respect to different data groups. Table 8-1 gives an

overview of organization and papers, which provide additional references, properties and physical characteristics, data bases for on-line data search. Changes in the contents and of data bases are reported by periodicals listed in Table 8-2. If no physical characteristics are available and self-experimental determination is not possible, the design engineer must rely on estimates and approximations. For such cases literature for the calculation of substance properties of pure fluid substances and mixtures are given in [8.1-8.3, 8.168.181.

534

8 Documentation and Calculation of Physical Characteristics

Table 8-1. Overview of organizations and papers which provide references, properties and physical characteristics and data bases [8.15]. Organization

List of data bases ~

Center of Information and Numerical Data Analysis and Synthesis (CINDAS) Purdue Industrial Research Park 2595 Yeager Road West Lafayette, Indiana 47 906 USA Chemie Information und Dokumentation Berlin (CIDB) Steinplatz 2 10623 Berlin Germany DECHEMA Stoffdatendienst Theodor-Heuss-Allee 25 60486 Frankfurt/M. Germany Fachinformationszentrum Energie, Physik, Mathematik GmbH, Karlsruhe 76344 Eggenstein-Leopoldshafen Germany Fachinformationszentrum Technik e.V. Ostbahnhofstr. 13/15 60314 Frankfurt/M. Germany (For further information: Gesellschaft fur Information und Dokumentation, Herriotstr. 5 60 528 Frankfurt/M., Germany) Information Analysis and Documentation Centre, Institut Francais du Petrole B.P. 311 92506 Rueil-Malmaison France Physical Property Data Service Institution of Chemical Engineers 164/171 Railway Terrace Rugby CV21 3HQ United Kingdom Technische Informationsbibliothek Welfengarten 1 B 30 167 Hannover Germany

L.M. ROSE: “Datenbanken fur physikalische Eigenschaften von Stoffen”, AchemaJahrbuch 1977/79, pp. 32-35. List of 22 data bases to calculate physical and chemical substance properties. EUSIDlC Database Guide 1981, A. TOMEERG: Learned Information Publisher, New York. ISBN 0904933 13 X. List of more than 1100 bibliographic data and numerical data bases of all fields including online access within Europe. How to find Chemical R.E. MAIZELL: Information, J. Wiley & Sons, New York 1979. ISBN 0-471 -56531 -8. References of more than 64 data centres for the chemical industry in the fields of energy, environment, physics, chemistry, materials. H. BEHRENS and G. EBEL:Physikdaten Datensamrnlung in der Physik. Zentralstelle fur Atomkernenergie-Dokumentation, 76344 Eggenstein-Leopoldshafen, Kernforschungsstelle, Germany. Information of more than 2000 data collecting systems. W. FRATZSCHER, H.P. PICHTand H.J. BIFFRICH:“Bemerkungen zur Ermittlung und Bereitstellung von Stoffdaten fluider Systeme fur verfahrenstechnische Berechnungen”, Chem. Techn. 29 (1977) NO. 7, pp. 361-420. List of 161 reference books and tables in the field of chemical engineering. J. HILSENRATH: “Summary of On-line or Interactive Physico-Chemical Numerical Data Systems”, NBS technical note 1122. Gov. Printing Office, Washington D.C. Description of 51 interactive systems to identify substances in a spectrogram, preparation of thermodynamic and transport properties of pure substances and mixtures, presentation of alloy properties, generation of thermodynamic tables, process simulation, process design. Encyclopedia of Information Systems and Services, 4th ed., Gale Research Company, Detroit (Mich.). Worldwide overview of EDP-data bases, creator and supplier; data network including operator; data collections, library services, data centres.

8 Documentation and Calculation of Physical Characteristics

535

Table 8-2. Regular periodicals providing information of new data bases [8.15]. Periodicals

Publisher

ISSN No.

CODATA Bulletin

Pergamon, Oxford

0366-151X

CODATA Newsletter

Dreyfus, Bertrand & Cie, Paris

0538-6918

Journal of Chemical and Engineering Data

American Chemical Society Books and Journals Division, Washington

0021 -9568

Journal of Chemical Information and Computer Sciences

American Chemical Society Books and Journals Division, Washington

0095-2338

Journal of Chemical Thermodynamics

Academic Press Inc., London

0021 -9614

Journal of Physical and Chemical Reference Data

American Chemical Society and American Institute of Physics for the National Bureau of Standards

0047- 2689

Online Review

Learned Information, Oxford, New York

0309-314X

536

8 Documentation and Calculation of Physical Characteristics

References [8.10]

LANDOLT-BORNSTEIN (HRsG.): Zahlenwerte und Funktionen. Springer, Berlin from 1961. PERRY,J. H. : Chemical Engineers Handbook. McGraw-Hill Book Comp., New York 1973. VDI-Wurmeatlas. Deutscher Ingenieurverlag, Dusseldorf 1975. WASHBURN, E. W. : International Critical Tables of Numerical Data, Physics, Chemistry and Technology. McGraw-Hill Book Comp., New York, from 1933. GROSSE, L. : Arbeitsmappe fur Mineralolingenieure. Deutscher Ingenieurverlag, Dusseldorf 1951. ~ M M E R M A N S ,J . : Physico-Chemical Constants of Pure Organic Compounds. Elsevier Publishing. New York 1950. LANCE,N. A. : Handbook of Chemistry. Handbook Publishers, Sundusky, Ohio

[8.11] [8.12] [8.13] [8.14] [8.15] [8.16]

[8.17]

1952.

D'ANs, J., and LAX,E. : Taschenbuchfur Chemiker und Physiker. Springer, Berlin 1949.

HODGMAN, C. D., WEAST,R. C., and SELBY,S. M : Handbook of Chemistry

[8.18]

and Physics. Chemical Rubber Publishing, Cleveland 1962. NIKOLSKI, B. P.: Handbuch des Chemikers, 3 Vols. Verlag Technik 1959. STAUDE, H. : Physikalisch-chemisches Taschenbuch, 2 Vols. Geest und Portig, Leipzig 1949. v. VOGEL,U. : Chemiker-Kalender. Springer, Berlin 1956. LEHRSTUHL FUR BCHNISCHE CHEMIEA, UNIVERSITAT DORTMUND: Das Aufsuchen von Stoffwerten. Dortmund 1970. CODATA: Internationaf Compendium of Numerical Data Projects. Springer-Verlag, Berlin 1969. SPRINGE,W., and K R ~ G E R H.: , Chem. Ing. Tech. 54 (1982) 4, 363-368. REID,R. C., PRAUSNITZ, J. M., and SHERWOOD,T. K. : Properties of Gases and Liquids. McGraw-Hill Book Comp., New York 1977. Autorenkollektiv: Berechnung thermodynamischer Stoffwerte von Gasen undFlussigkeiten. Deutscher Verlag fur Grundstoffindustrie, Leipzig 1966. WESTMAIER, S., et al. : Verfahrenstechnische Berechnungsmethoden, part 7: Stoffwerte. VCH, Weinheim 1985.

Thermal Separation Processes: Principles and Design Klaus Sattler, Hans Jacob Feindt copyright 0VCH

Verlagsgesellschaft m h H , 1995

General References * Handbooks, comprehensive reference works, abstracts

[0.1] BARTHOLOME, E., BIEKERT,E., HELLMANN, H., LAY, H., WEIGERT,H., and WEISE,E. (eds.): UflmannsEncyklopadie der technischen Chemie. Vol. 1 Grundlagen; Vol. 2 Verfahrenstechnik I (Grundoperationen); Vol. 3 Verfahrenstechnik I1 und Reaktionsapparate. 4th Ed. Verlag Chemie, Weinheim 1973. [0.2] Autorenkollektiv: Lehrbuch der chemischen Verfahrenstechnik. Deutscher Verlag fur Grundstoffindustrie, Leipzig 1983. [0.3] ADOLPHI, G., and ADOLPHI, H. V.: Grundziige der Verfahrenstechnik. Deutscher Verlag fur Grundstoffindustrie, Leipzig 1970. [0.4] GRASSMANN, P., and WIDMER, F. : Einfiihrung in die thermische Erfahrenstechnik. Walter de Gruyter & Co., Berlin 1974. [0.5] KASSATKIN, A. G. : Chemische Erfahrenstechnik, 2 Vols. Deutscher Verlag fur Grundstoffindustrie, Leipzig 1966. [0.6] VAUCK,W. R. A., and MULLER,H. A.: Grundoperationen chemischer Verfahrenstechnik. Verlag Chemie, Weinheim 1982. [0.7] EUCKEN,A., and JAKOB,M.: Der Chemie-Zngenieur, 12 Vols. Geest und Portig, Leipzig 1940. [0.8] T~EYBAL, R. E.: Mass Transfer Operations. McGraw-Hill Book Comp., New York 1968. [0.9] COULSON, J. M., and RICHARDSON, J. F. : Chemical Engineering. Pergamon Press, London 1964. W. L., and BANCHERO, J. T.: Zn[O.lO] BADGER, troduction to Chemical Engineering. McGraw-Hill Book Comp., New York 1955.

*

[0.11] PRATT,H. R. C.: Countercurrent Separation Processes. Elsevier Publishing Comp., Amsterdam 1967. [O. 121 SMITH,B. D. : Design of Equilibrium Stage Processes. McGraw-Hill Book Comp., New York 1963. [0.13] SAWISTOWSKI, H., and SMITH,W.: Mass Transfer Process Calculations. Interscience Publishers, New York 1963. [0.14] PAWJDW,K. F.: Beispiele und Ubungsaufgaben zur chemischen Verfahrenstechnik. Deutscher Verlag fur Grundstoffindustrie, Leipzig 1966. [0.15] CLARKE,L., and DAVIDSON, R. L.: Manual for Process Engineering Calculafions. McGraw-Hill Book Comp., New York 1962. [0.16] MCCABE,W. L., and SMITHJ. C.: Unit Operations of Chemical Engineering. McGraw-Hill Book Comp., New York 1967. [O. 171 PERRY, J. H. : Chemical Engineers' Handbook. McGraw-Hill Book Comp., New York 1973. [O.I8] CREMER,W., and DAVIES,T.: Chemicaf Engineering Practice, 10 Vols. Butterworth Scientific Publications, London, from 1955. [0.19] KIRK, R. E., and OTHMER,D. F.: Encyclopedia of Chemical Technology, 3rd Ed., J. Wiley and Sons, New York, from 1978. [0.20] ONKEN, U. : Thermische Erfahrenstechnik. Carl Hanser, Munchen 1975. [0.21] DECHEMA-Monographien. Verlag Chemie, Weinheim. [0.22] MIESSNER,H. : Fortschritte der Verfahrenstechnik. Verlag Chemie, Weinheim.

The reader is directed to the end of each chapter for specific references within each chapter.

538

General References

Deutscher Verlag fur Grundstoffindu[0.23] DREW, T. B., HOOPES,J. W., and VERstrie, Leipzig 1984. T. (eds.): Advances in Chemical MEULEN, [0.28] GERHARTZ,W., et al. (eds.): “Unit Engineering. Academic Press, London. Operations”. Ullmann’s Encyclopedia of [0.24] BAYERAG (ed.): Verfahrenstechnische Industrial Chemistry. Vols. B2, B3. VCH Berichte. Verlagsgesellschaft, Weinheim 1988. [0.2 51 DECHEMA : DECHEMA-Literatur-Schnell[0.29] COULSON,J. M., RICHARDSON,J. F., dienst. BACKHURST, J. R., and HARKER,J. H.: [0.26] MERSMANN, A. : Thermische Verfahrens“Particle Technology and Separation technik. Springer-Verlag, Berlin, HeidelProcesses”. Chemical Engineering. berg 1980. Vol. 2. Pergamon Press, Elmsford, [0.27] WEISS,S., und MILITZER, K.-E.: ThermiOxford 1991. sche Wrfahrenstechnik I und II. VEB

Thermal Separation Processes: Principles and Design Klaus Sattler, Hans Jacob Feindt copyright 0VCH

Verlagsgesellschaft m h H , 1995

Index

Absorbent 243 Absorber 262 -, absorption reactor 267 -, bubble column 264 -, column 268 -, film 268 -, washer, scrubber 268 Absorption 239 Absorption column 248-262, 268 Absorption equilibrium 44-51 Absorption machines 272 Absorption regeneration, desorption 239, 263-277 -, cocurrent 248 -, countercurrent 248 -, equilibrium 44-51 -, isotherm 46 Activity coefficient 16 Adsorbent 291 Adsorbent load, line of constant 56 Adsorbent properties 295 Adsorbent regeneration, desorption 311-315 Adsorbers 310-31 1 -, fixed bed 311 -, fluid bed 312 -, moving layer 313 Adsorption 281 -316 -, countercurrent 308 -, hysteresis 56, 58 -, isobar 55 -, isostere 55 -, isotherm 55 -, kinetics 293-301 -, multistage crosscurrent 307 -, single stage 301 Analogy, heat and mass transfer 73 Antoine equation 30 ARD-extractor (asymmetric rotating disc contactor) 446 Axial mixing 416 Azeotrope 41 Azeotrope rectification 125 Azeotropic point 41

Backflow model 417 Balance 8-13 -, energy 12 -, enthalpy 12 -, exergy 12 -, heat 12 -, mass 10-12 Belt dryer 364 Binary distillation 110 Binary system 37 Binodal curve 24 Boiling curve 38 Boiling diagram 38 Boiling point elevation 51 Boiling point maximum 40 Boiling point minimum 40 Bollmann extractor 464 Breakthrough curve 300 Brunauer, Emmet, Teller equation 59 Bubble column 264 Bubble point line 39 Bunsen’s absorption coefficient 46 Capillary condensation 56 Carrier distillation 113- 116 Cascade extractor 464 Centrifugal extractor 441, 453 Chamber dryer 363 Channeling 214 Chemical equilibrium 48 Chemical plant 1 Chemical potential 15 Chemical precipitation 477 Chemisorption 48-51, 239, 263, 281 Circulation evaporator 501 Clausius-Clapeyron equation 18, 29, 53 Cleaning of flue gas 287 -, Lurgi-Kontisorbon process 287 -, Lurgi-Supersorbon process 285 -, Walther process 240 - , Wellmann-Lord process 263 Climbing film evaporator 503 Coadsorption 301 Cocurrent absorption 248

540

Index

Cocurrent distillation 222 Cocurrent operations 77 Cocurrent principle 3 Column accessories 218 -, hold-down plate 221 -, liquid distributor 224-229 -, support plate 221 Column diameter 248 -, absorption 248 -, extraction 419 -, rectification 164 Column internals 165 -, random packing 196 -, stirring device 446 -, structured packing 203-205 -, tray 167 Column with stirring device 446 Conjugation line 25 Consistency test 36 Contact drum dryer 387 Contact drying 340-351 Contact-mixing dryer 381 Control 220 -, of a dryer 387 -, of a rectification column 220 Convective drying 331, 340-349 -, gas and heat requirement 340-343 -, mass and heat transfer 331 -, methods 336 Cooling requirement -, absorption 247 -, adsorption 304 -, crystallization 506 Countercurrent absorption 248 Countercurrent adsorption 308 Countercurrent crystallization 520, 524 Countercurrent direct condensation 233 Countercurrent distillation 119 Countercurrent extraction 407 Countercurrent operation 79 Countercurrent principle 3 Countercurrent sublimation 530 Cross section -, absorber 248 -, adsorber 303 -, crystallizer 516 -, dryer 358 -, extractor 419 -, rectification column 164 Crosscurrent adsorption 307 Crosscurrent extraction 403 Crosscurrent operation 94 Crystal growth 508

Crystal product rate 500 Crystallization 475 -, adductive 477 -, by cooling 484, 485 -, by evaporation 484 -, by vacuum cooling 484 -, column 521, 524 -, desublimation 475, 524-530 -, from a melt 481, 521-524 -, from a solution 475, 484, 500-508 -, heat, enthalpy 506 -, kinetics 508 -, methods 475 Crystallizer 511 -516 -, crystallization from solution 511 -, melt crystallization 523 -, sublimator 529 Dalton’s law 38 Data bases 533 Design procedure of separation plant 94 Design, sizing 248 -, absorber 248 -, adsorber 303, 307 -, crystallizer 516 -, distillation device 106, 109, 147 -, dryer 357 -, evaporator 490 -, extractor 409, 419, 424, 456 Desorption 239, 263-277, 282, 311-315 Desublimation 524 Dew line 39 Diameter 248 -, absorption column 248 -, adsorber 303 -, crystallizer 516 -, dryer 358 -, extractor 419 - , rectification column 164 -, still 106 Dielectric drying 352 Diffusion 68-70 - , coefficient 71 -, model 418 Diffusional distillation 131 Dimensionless numbers 73 Dispersed phase 407 Displacement desorption 314 Distillation 101 -, batch 103 - , cocurrent 222 -, column 165

Index

-, continuous 107

-, countercurrent 102, 119-216 -, equipment 107

-, evaporator 500 -, fractional 104 -, still 107

Distillation operation, processes 101 -, azeotrope 125 -, carrier 113-116 -, diffusional 131 -, extractive 127 -, flash 111-113 -, heteroazeotrope 129 -, molecular 116 -, rectification 119 -, two pressures 130 Distribution coefficient 19 Double cone dryer 385 Dryer 357-389 Drying 317 -, contact 349-351 -, convective 340-349 -, dielectric 352 -, freeze 355-357 -, principles 319 -, radiative 351 Drying kinetic 335-340 Drying rate 335-340 Drying time 335, 337 Drying with fresh and recycled air 347 Dubinin isotherm 59 EC-extractor (enhanced coalescence extractor) 448 Energy balance 12 Enhancement-factor 257 Enrichment section 134, 166, 422 Enthalpy 14 -, Gibbs’ free 14 -, latent heat of phase change 65 Enthalpy concentration diagram 110 Enthalpy, Mollier’s diagram (moisture loading diagram) 327 Entrainer sublimation 525 Entropy 14 -, of mixing 32, 67 Equilibrium, conditions of 16 Equilibrium diagram 38 Eutectic system 60, 64 Eutectic point 64 Evaporator 500, 501, 504 -, circulation 501 -, continuous, one pass 505

-, with stirring gear 504 Excess enthalpy 32 Exergy balance 12 Extract phase 393 Extraction 393 -, high pressure 393, 463-470 -, leaching, solid-liquid 393, 458-463 -, solvent, liquid-liquid 395-399 -, solvent selection, criteria 399 Extraction column 426-456 -, random, irregular packing 430 -, sieve tray (perforated plate) 430 -, spray 430 -, structured packing 430 Extraction equilibrium 19 Extraction factor 402, 406, 410 Extractive distillation 42 Extractive rectification 127 Extractor 424-456 -, agitated column 424-456 -, cascade 464 -, centrifugal 441, 453 -, column 426 -, column with stirring devices 440 -, leaching 464 -, mixer-settler 425 Falling film evaporator 505 Feed stage 157 -, optimum 156 -, selection 156 Fenske, Underwood, Gilliland method 155 Fick’s law 69 Film absorber 268 -, dryer 381 -, evaporator 503 Fixed bed adsorber 311 Flash distillation 111 Flash evaporation 498 Flooding point 420 -, extraction 420 -, rectification 202-209 Flue gas treatment -, Walther process 240 -, Wellman-Lord process 263 Fluid bed, fluidized bed 311 -, adsorber 311 -, dryer 366 Fluidization velocity, minimum 368 Fractional crystallization 520 Fractional distillation 104 Free void space 199-201, 203, 295

541

542

Index

Freeze concentration 478, 519-520 Freeze drying 355-357 Freundlich isotherm 59 Froth section 168 Fugacity 30, 32 Gibbs-Duhem equation 15, 36 Gibbs-Kelvin equation 56 Gibbs’ phase rule 18 Graesser contactor 450 Heat and mass transfer 69, 73 -, convective drying 331 -, fixed bed adsorber 305 Heat balance 8, 12 Heat of mixing 32, 67 Heat of vaporization 29 Heat pump 139, 141 -, drying 336 -, rating 144 -, rectification 142 Heat requirement, consumption -, crystallization 507 -, drying 340, 350 -, rectification 136, 161 -, simple distillation 109 -, solution concentration 486 Heat transfer 74 -, crystallization 507 -, distillation 108 -, drying 360 -, evaporation 490 Heat transformer 142 Henry’s law 44 Heteroazeotrope rectification 129 HETS (height equivalent to one theoretical stage) 86, 214, 254, 412 High pressure extraction 393, 463-470 Hildebrandt extractor 464 Holdup 215 -, extraction 420 -, rectification 215 HTU, NTU method 91 -, absorption 255 -, adsorption 309 -, extraction 414 Humid gas 324 -, change of state in convection drying 334-335 -, properties 324-331 Hunter-Nash method 411 Hypersorption 310

Ideal mixtures 37 Internal energy 14 -, Gibbs free 14 Karr column Kinetic theory separation Kiihni column

443 of countercurrent 90 447

Langmuir isotherm 59 Latent heat 65 Leaching 393, 458-463 Lever rule of phase mixture 24 Liquid phase holdup 215, 419 Load limit 202, 211 Local tray efficiency 190 Lowering of freezing point 63 LUB theory 303 Lurgi-Westfalia extractor 454 Maldistribution 214 Mass balance 10-12 Mass transfer 68-77 -, convection 72 -, diffusion 69 Mass transfer coefficient 72 -, film model 75 -, in packed beds 258 Mass transfer zone 300 Material balance 8 McCabe-Thiele diagram 82, 149, 163 Melting pressure curve 62 Merkel, Ponchon diagram 110 Minimum work of separation 67 Mixer-settler 425 Mixture types 38-44 - , heteroazeotropic 41 -, ideal 37 -, real 37 -, with boiling point maximum 40 -, with boiling point minimum 40 Molecular diffusion 69 Molecular distillation 116 Mollier’s h,X-diagram 327 Multistage drying 348 Murphree stage efficiency 6, 84 Nernst’s distribution law 19, 22 Nonadiabatic rectification 222 NTU, HTU method 91 Nucleation 509

Index Oldshue-Rushton extractor 447 Operating diagram 249 -, absorption 249, 251 -, adsorption 308, 309 -, cocurrent operation 77 -, countercurrent operation 79 -, crosscurrent operation 94 -, extraction 402, 404, 406, 409, 411, 423 -, rectification 151 Operating mode 7 -, alternating (cyclic) 8 -, batch 8 -, cocurrent operation 3 -, continuous 8 -, countercurrent operation 3 -, crosscurrent operation 6 -, cyclic, alternating 8 Ostwald’s absorption coefficient 46 Overall mass transfer 74 Overall mass transfer coefficient 76 Packed column 196 Packing 196, 203 Paddle dryer 381 Parallel flow, multiple effect evaporation 489 Partial condensation 230 Partial condenser 231 Partial molar volume 18 Particle diffusion 299 Penetration theory 77 Phase 14 -, continuous 14 -, dispersed 14, 407 Phase diagram 29, 64, 65 Phase equilibrium 5, 14-67 -, gas-solid 52-60 -, liquid-liquid 19-27 -, liquid-solid 60 -, vapor-liquid 28, 37, 44 Phase flow 3-6 -, cocurrent 3, 78 -, countercurrent 3, 79 -, crosscurrent 6, 94 Phase interface area, interface 75, 91, 261, 421 Phase rule 18 Physical sorption 239 Physical, physico-chemical properties 533 Plate dryer 364 Pneumatic dryer 374 Podbielniak extractor 455 Ponchon-Savarit method 155 Pore diffusion 299

Porosity 199-201, 203, 295 Precipitation 477 Pressure diagram 38 Pressure drop 195 -, column tray 195 -, fixed bed, packed bed 304, 372 -, fluidized bed 372 -, packed column 210 Pressure swing adsorption 282 Production plant 1 Pulsation 434, 438 Pulse sieve tray column 438

QVF extractor

448

Radiative drying 351 Raffinate 393 Random packing 196 Range of operation 184 -, column tray 184 -, packed column 202 Raoult’s law 37-44, 51 -, mixture 37 -, solution 51 Rating factors, heat pump 144 Rayleigh’s equation 104 Rectification 102, 119- 130 -, adiabatic 119 -, batch 158 -, continuous 134 -, nonadiabatic 222 -, semi-continuous 163 Reduction of energy consumption, energy savings 138 -, drying 343 -, rectification 138 Reflux 79, 136 Reflux ratio 136, 157 -, minimum 157 -, optimum 157 Regeneration -, of absorbent 263 -, of adsorbent 311 -, of solvent 393 Relative humidity 5 5 , 324 Relative volatility 67 Residence time 6 -, mean 6 Revolving extractor 465 Robatel extractor 455 Roller dryer, drum dryer 381 Rotating disc contactor 446 Rotating drum dryer, rotary dryer 364

543

544

Index

Saddles 199, 200 Sauter diameter 421 Scheibel column 447 Separation factor 67 Separation processes, operation 4 -, mechanical 4 -, thermal 3 Separation spray crystallization 478 Separation stage 5 -, practical 5 -, theoretical 5 SHE extractor 449 Solubility diagram 62 Solubility of gases 48 Solubility of solids 60 Solution concentration by evaporation 484, 485-500 -, multistage 487 -, multistage flash 498 -, single stage 486 Solvent -, absorption 243 -, extraction 395-399 Solvent extraction 395-399 -, countercurrent 407 - , countercurrent distribution 424 -, countercurrent with extract reflux 421 -, crosscurrent 403 -, differential batch 403 -, multistage 403-407 -, single stage 400-403 Solvent ratio -, absorption 248, 250 -, extraction 402, 406, 409 Solvent requirement -, absorption 243 -, extraction 408 Sorbent, auxiliary component 243 -, absorption 243 -, adsorption 291 -, azeotropic rectification 124 -, extraction 399 -, extractive rectification 124 Source, sink 10 Specific surface 297 -, of adsorbent 297 -, of packing 199-201, 203 Spray crystallization 478 Spray dryer 377 Stage 5 -, practical 5 -, theoretical 5

Stage efficiency 5-6 Stage method 82-89 -, absorption 250-262 -, adsorption 309 -, distillation 147-157 -, extraction 404, 407, 459, 462 -, stage-to-stage calculation 413 Stage model, cell model 417 Steam distillation 112, 115 Steam requirement, consumption 490 -, solution evaporation 490 Stefan flux 69 Still 107 Stripping 239, 263, 274 Stripping section 134, 422 Sublimation 52, 524-530 -, apparatus 529 -, entrainer 525 -, fractional 530 -, single stage 524 -, vacuum 525 Sublimation pressure curve 53 Super-saturation 62 -, allowed 62, 510 -, curve 62 Surface drying 332 Theoretical stage 82 Theory of separation stages 79-82 Thermodynamic system 14 Tie line 24 Transfer unit 92 -, height 92 -, number of 89 Tray 167-196 -, bubble-cap 174 -, channel 175 -, cross-flow 176, 188 -, mixing 182 -, performance comparison 183 -, sieve 172 -, valve 177 Tray column 165 Tray efficiency 190 -, absorption 254 -, extraction 412 -, rectification 190 Triangular diagram 24 Triple-point 29 Tumbler dryer 385 Tunnel dryer 364 Turbulence theory 77 Two layers theory 419

Index Two phase range 182 Two pressure distillation 130 Vacuum sublimation 525 Valve tray 177 Vapor compression 115 -, distillation 115, 142 -, mechanical 142, 492 -, solution evaporation 492, 495 -, steam jet 496 Vapor pressure curve 29 Vapor pressure lowering 51 Wall effect on flow 214 Washer, scrubber 269 -, jet 271 -, radial flow 270 -, ring slot 270

-, rotary 273

-, rotation column 272 -, spray 269

-, Stroder 272 -, Theisen 273 -, Venturi 271

Waste gas treatment 240, 263 -, Lurgi-Kontisorbon process 287 -, Lurgi-Supersorbon process 285 Weir on a plate 178 Wet product 320 -, properties 320 -, transfer of humidity 325 Wet-bulb temperature, cooling limit 3 34 Whitman’s two film theory 75 yx diagram 23

545

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Batch operating rectification unit DN 30 Ibi-separation of liquid inixturcs (maximurn theoretical pla~es:approx. 60) for opcrations tinder normal pressure and under vacuum equipped with an electromagncticall~~ controlled liquid divider-column head with inclined condenser and an elcctroinafnetically controlled side takc-otf’ device for gaining in-hetween fractions Process engineering apparatus and plants tor the laboratory

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Sulzer Chemtech Ltd, CH-8404 Winterthur, Switzerland, Telephone 052-262 50 28/262 45 17, Fax 052-262 00 67 Great Britain: Fax 01252-51 86 36; Venezuela: Fax 02-93 88 77; Singapore. Fax 065-861 15 16, India: Fax 0212-33 86 06; Netherlands: Fax 03440-205 38; USA, Canada: Fax 0416-213 10 31; Japan: Sumitomo, Fax 03-3233 96 26 In addition, there are Sulzer offices and representations in more than 100 countries

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Laboratory and Pilot Plants for Distillation, Absorption, Adsorption, Extraction and Reaction

FiSSCHER

Micro-SPALTROHRTM-columns for the gentle distillative separation of small charges. Column with the lowest pressure loss and the highest separation efficiency.

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RALU-PAK 250 YC the right solution for many applications RALU-PAK 2 5 0 YC is a modern, uniform mass-transfer packing with extraordinary separation capacity within a wide operation range, which also displays an extremely low pressure drop. RALU-PAK 2 5 0 Y C can be employed for destillation, rectification, absorption and desorption. RALU-PAK 250 YC is the ideal solution wherever large gas volumes and/or small volumes of liquid are to be handled. Together with the system liquid distributors, liquid collectors and support grids which were specially designed for system packings, RASCHIG offers a solution concept which has proven its worth throughout the world. We would be glad to provide you with additional information.

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MESSO-

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No matter whether complete turnkey plants or systems for individual Rrocess s t e m are concerned - projects (ike this call for experienced specialists We supply complete custom-tailored solutions for the 0 Crystallization of over 100 inor

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UMBRELLA-TYPEBUBBLE CAP Instead of columns with structured packings bubble cap trays are recommended if - there is a danger of strong contamination, -there are varying liquid loads in the column, - a reflux ratio is required which is below the minimum liquid load of the package or - many feeds and sidestreams have to be realized

Standard SCHMlDDlNG bubble cao

If you are interested in getting more information and more details about umbrella-type bubble caps please phone us or send us a fax SCHMlDDlNG umbrella-type bubble cap tray DN 1800

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Emdener Str lb D - 50735 Cologne (Germany)

Telephone +49 221 7174 - 01 Telefax +49 221 7174 - 234

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KClHNl

AP 3 (ALL-PHASES)

DISCOTHERM B

I I X I N G A N D KNCADING

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L O S t D DESIGN 'ROCESSING UNDER PRESSURE i N D i O R VACUUM O N I K O 1 I AH1 1 KkSIDFNCF TIMF

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Column Internals for Packed Columns ~

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High efficiency dumped packings with low pressurc drop, great operating range. made of inetnls, ceramics a n d plastics; tower pnckings such as pallrings, DINPAC, ENVIPAC, R-PAC, SR-PAC 2nd rolling strtictiiral grid packing for rectification, absorption and liquid/liquid-extraction processes. Droplet \eparators, innde trom plastics and metals. Liquid distributors, liqtiid redistributors. supports made from metals and plastics. Noules and noule
ENVICON Engineering GmbH offer: PC-software FDPAK, HTUPAK and TRAYS for packed and tray columns design

The PC-program FDPAK was evaluated according to the SRD-model. This program allows the cnlculation of hydraulic parameters: flooding gas velocity, hold-up and pressure drop at flooding point and under operating condition, t'or about 200 dumped packings and structurnted packings,. - The progrm HTUPAK can be used for design of packed bed height of various pnckings for absorption. stripping, purification and rectification o l groundwater. - TRAYS allows the design of sieve-trays, sieve-valve, tunnel-valve. Baycr-valvcs. bubble trays, Vario-Flex-valve - including all hydraulics parameters and tray efficiency for absorption, desorption and rectification. Thc programs FDPAK, HTUPAK and TRAYS were worked out on the base of the rcsults of theoretical and cxpcrimentnl papers. ENVICON Engineering offer fast experimental evaluations of model parameters for packed and tray columns with own pilot-plant equipment lor rectification. absorption and liquid/liquid-extraction. -

~

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Adress for correspondence: i w d d h o i 2-(1 . 1)-a>i37i h d i h c ~ 1.1ldct,ll1. I ) I I ~ ' ~ / ~ I I ( .~I~ICI;IX~ / ~ ~ I~I -I (~ ~~ ( ~ / ~ O ( , ~ / ~ ~ ~ - ~ X

You can purify organics

10ppm

levels using Sulzer C h e m ~

Or, i f you need tonnage samples, you can rent our mobile

lech's Fractional Crystallization Process Without solvents

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arid without ciysial sluriies More thari twenty industiial

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Sulzer Cherntech designs and

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SUUER

WCHEMTECH

Fractional Crystallization

Sulzer Chemtech Ltd, PO Box, CH-9471 Buchs, Switzerland, Phone [ 41 181-75603 11, Fax [ 41 181-7564012

5