2064772 - university of wales, newport
TRANSCRIPT
University of South Wales
2064772
INTEGRATED DESIGN OF CHEMICAL PLANTS
WITH ENERGY CONSERVATION (THE DESIGN
OF AN ENERGY EFFICIENT STYRENE PLANT)
BY
AUDAY ESMAIL SAEED
THESIS SUBMITTED TO THE C.N.A.A. IN
PARTIAL FULFILMENT IN CANDIDATURE FOR
THE DEGREE OF DOCTOR OF PHILOSOPHY
DEPARTMENT OF SCIENCE AND CHEMICAL ENGINEERING
THE POLYTECHNIC OF WALES,
TREFOREST, PONTYPRIDD,
MID GLAMORGAN,
CF37 1DL
IN COLLABORATION WITH BP, LONDON
OCTOBER, 1990
DEDICATION
TO MY DEAREST BELOVED (JACQUELINE)
DEDICATED TO:
MY FATHER AND MOTHER
AND
MY SISTERS: SAHAR AND RAGHAD
MY BROTHERS: HAKI, KUSAY, SA'AD, AND RA'AD
MY SISTER IN LAW: ENTESAR
MY NIECE: NUR
CONTENTS
Page No.
Acknowledgements i
Declaration ii
Certificate of Research iii
Abstract iv
Nomenclature v
CHAPTER ONE : Introduction 1
1.1 Chemical Plants and Energy Requirements 2
1.2 Thesis Layout 3
CHAPTER TWO : On the Process Design and Energy
Integration 6
2.1 Process Synthesis 8
2.1.1 Synthesis With a Feasible Flowsheet
in Hand 8
2.1.2 Synthesis Without Having an Initial
Flowsheet 9
2.1.2.1 Heuristic Approach 9
2.1.2.2 Thermodynamic Approach 10
2.2 Flowsheet Synthesis Problem 10
2.2.1 Heat Exchanger Network Synthesis 11
2.2.1.1 Heat Exchanger Network
Specification 13
2.2.1.2 Different Approaches For Minimum
Utility Usage 13
2.2.1.3 Heat Exchanger Network
Representation 14
2.2.1.4 Minimum Utility Targets and the
Pinch Concept 18
2.2.1.4.1 Graphical Method 19
2.2.1.4.2 Problem Table Method 25
2.2.1.5 Appropriate Placement of Utilities 29
2.2.1.6 Grand Composite Curve 32
2.2.1.7 Constraints 34
2.2.1.7.1 Heat Capacity Flowrate
Constraint 34
2.2.1.7.2 Number of Process Streams
Constraint 36
2.2.1.8 Stream Splitting Technique 37
2.2.1.9 Minimum Number of Units 38
2.2.1.10 Utility Pinches 39
2.2.1.11 Threshold Problem 39
2.2.1.12 Process Improvement and
Modification. 41
2.2.2 Separation System Synthesis 43
2.2.2.1 Energy Integrated Distillation
Column 48
2.2.2.2 Appropriate Placement of
Distillation Column 49
2.2.3 Heat and Power System Synthesis 53
2.2.3.1 The Integration of Heat and Power
in the Process Network 54
2.2.3.1.1 Appropriate Placement of Heat
engines 54
2.2.3.1.2 Appropriate Placement of Heat
Pumps 55
2.2.3.2 Selection of the Right Heat Engine 58
CHAPTER THREE : Mass and Energy Balance 60
3.1 Mass Balance 60
3.1.1 Overall Mass Balance 61
3.1.2 Alkylation Process Step 63
3.1.2.1 Mass Balance Over the Alkylator 64
3.1.2.2 Mass Balance Over the Distillation
process 66
3.1.3 Dehydrogenation Process Step 66
3.1.3.1 Mass Balance Over the Dehydrogenator 69
3.1.3.2 Mass Balance Over the Distillation
Process 69
3.2 Energy Balance 72
3.2.1 Alkylation Process Step 72
3.2.2 Dehydrogenation Process step 74
3.3 Concluding Remarks 75
CHAPTER FOUR : The Selection of an Optimum
Unintegrated Distillation Sequence 78
4.1 Heuristics Used 78
4.2 Alkylation Process 79
4.2.1 Identification of Possible Unintegrated
Sequences 79
4.2.2 Heuristics Application 83
4.2.3 Heuristics Philosophy 85
4.2.4 Minimum Reflux Ratio Calculations 87
4.2.5 Minimum Number of Plates Calculations 89
4.2.6 Optimum Reflux Ratio 89
4.2.7 Energy Consumption 91
4.2.8 Concluding Remarks 101
4.3 Dehydrogenation Process 102
4.3.1 Concluding Remarks 104
CHAPTER FIVE : Energy Integration in the Styrene
Process 106
5.1 Alkylation Process 106
5.1.1 Streams Extraction 106
5.1.2 Targetting and Related Design 110
5.1.3 Energy Saving Techniques and Process
Improvement 116
5.1.3.1 Energy Recovery By Inspection 116
5.1.3.1.1 Analysis of the Network Design 121
5.1.3.2 Increasing Energy Recovery By Process
Improvement 130
5.1.3.2.1 Process Examination For More
Improvement 137
5.1.4 Process Utility Levels 141
5.1.5 Concluding Remarks 141
5.2 Dehydrogenation Process 144
5.2.1 Streams Extraction 144
5.2.2 Targetting and the Old Design Failures 144
5.2.2.1 The Design That Reaches the Energy
Targets 150
5.2.3 Process Improvements 153
5.2.3.1 Further Energy Recovery by
Utilizing Stream No.2 155
5.2.3.1.1 Tick Off Rule Application 160
5.2.3.2 The Reduction of Utility Consumption
By Utility Generation 163
5.2.3.2.1 Cold Utility Reduction By Steam
Generation 164
5.2.4 Dehydrogenation Process Utility Levels 170
5.2.5 Concluding Remarks 173
CHAPTER SIX : The Utility Interface With the Process
Design 174
6.1 Dehydrogenation Process and Utility
Consumption 174
6.2 The Use of Hot Utility and Stack Loss 175
6.3 Flue Gas in Process Context 179
6.3.1 The Network Design Analysis After
Introducing the Flue Gas as a Process
Stream I 83
6.3.1.1 The Utilization of the Heat Load on
Stream No.l 186
6.4 The Utilization of the Steam Raised in the
Process 187
6.5 Concluding Remarks 191
CHAPTER SEVEN : The Effect of Process Integration on
the Steam and Power System 194
7.1 Combined Heat and Power System in Styrene
Process 194
7.1.1 Application of a Combined Heat and Power
System in the Styrene Plant Before
Energy Integration 195
7.1.1.1 Power Generation in the Plant 197
7.1.1.2 Fuel Consumption in the Boiler House 197
7.1.2 Application of a Combined Heat and Power
System in the Styrene Plant After
Energy Integration 201
7.2 Concluding Remarks 203
CHAPTER EIGHT : Conclusions 205
8.1 Future Investigations 206
REFERENCES :
APPENDICES :
208
Appendix A : Mass and Energy Balance sampleCalculation 221
Appendix B : Vapour Pressure, Relative
Volatility and Average Relative
Volatility Data 226
Appendix C : Computer Program For Calculating
the Root of Underwoods Equations 237
Appendix D : The Data For Calculating the
Optimum Reflux Ratio 241
Appendix E : Flame Temperature Calculation 254
ACKNOWLEDGEMENTS
The author wishes to express his deep gratitude and
sincere thanks to his thesis supervisors Dr. M. S. Doulah and
Mr. F. B. Blakemore. Their continual help, active
encouragement, guidance, and invaluable suggestions and
discussions throughout the course of this research has made
me deeply indebted to them.
I would like to express my sincere thanks to Dr. G. Rees
in the Polytechnic of Wales for his help and support in
different research matters.
I gratefully acknowledge the collaboration provided by BP
Head Office, and I also extend my sincere thanks to Mr. J. B.
Rutter of BP, together with Gulf and BP for permitting me the
opportunity of visiting their plants.
It is a pleasure for me to acknowledge the assistance of
the computer centre in the Polytechnic of Wales for allowing
me to use and borrow some of their facilities.
Finally, I would like to thank all the colleagues who
shared the office with me (Neal, Philip, Amanda, Angela,
Chris, Carsten, Tanveer, and Johannes). They were such good
company and helped my days go by.
DECLARATION
Thisisto certify that neither this thesis nor any part of
it has been presented or is being concurrently submitted in
candidature for any other degrees.
Candidate
Dated '.October,
11
CERTIFICATION OF RESEARCH
This is to certify that, except where specific reference
is made, the work in this thesis is the result of the
investigation carried out by the candidate.
Candidate
ADirector of Studies
Dated: October,
ABSTRACT
Energy consumption is one of the main areas in the study of chemical process design. It is usually referred to as the critical element that is continuously needed for running a chemical process, and is daily effected by the prices of energy. Therefore, poor designs which are not energy integrated normally lead to less profit due to high consumption of energy. These simple economics are the reason for tackling the area of energy integration in process design. A styrene production process is taken to be the model process for carrying out the design work incorporating the various energy integration techniques.
A thorough review of the published work in this subject area was the first step in this research work. This has been followed by calculating mass and energy balances around the overall plant and the individual process steps, so that information about flowrates and energy consumed and released was obtained for the base case. After this all the possible distillation sequence configurations were tested in order to find the sequence that required least energy compared with all the other possible sequences. This step is the first part of integrating the distillation train. The second part considered the heat exchanger network associated with the distillation train and this has been taken in the context of overall process integration. "Pinch technology" was used as an aid for targeting the minimum hot and cold utilities required, designing the heat exchanger network that was compatible with the minimum use of utility and to seek further improvements on the process heat exchanger network which made it capable of recovering even more energy.
Utility supplies are designed with respect to the process design, hence the next step considered the interaction between the utility and process design. Thus, the utilities were introduced in a more efficient way, resulting in a better heat exchanger network and increasing the interprocess heat exchange. Finally the steam and power system in the styrene plant was tested in order to determine how much this system had benefited due to the overall efficiency of energy supply and demand.
IV
Nomenclature
Dimensions of the symbols are given in terms of mass (M),
Length (L), Time (T) and Temperature (0),
Symbol Definition Dimension
ANTA constant in the Antione equation ______
ANTE constant in the Antione equation ____
ANTC constant in the Antione equation ____
B moles or mass of the bottom stream
from distillation column per unit
time MT' 1
Bz benzene ____
Ci, Cz , Ca cold streams No.1, 2, 3
respectively ____
CB aromatic carbon __
Cp heat capacity flowrate ML2 T~ 2 Q- 1
fcp heat capacity L2 T~ 2 e- 1
(Cpa)nin minimum heat capacity flowrate of
branch a ML2 T- 2 6- *
(Cpb)nin minimum heat capacity flowrate of
branch b MI^T-^- 1
Cpc heat capacity flowrate of a cold
stream ML2 T- 2 6-*
Cph heat capacity flowrate of a hot
stream ML2 T- 2 6~ 1
D moles or mass of distillate per
unit time MT~ 1
DEB diethylbenzene ____
EB ethylbenzene ____
Et ethylene ____
F moles or mass of feed to MT~ 1
distillation column per unit time MT~ x
Hi, H2, Ha hot streams No.l, 2, 3
respectively _____
AHzs standard heat of reaction L2 T~ 2
&HF<prod.) enthalpy of formation for product L2 T~ 2
AHF(reac.) enthalpy of formation for reactant L2 T~ 2
&Hi enthalpy of component i L2 T~ 2
AHp enthalpy of products L2 T~ 2
&HR enthalpy of reactants L2 T~ 2
J number of components ____
L moles or mass of reflux to
distillation column per unit time ____
m molar or mass flowrate MT~ 1
N total number of streams including
utilities ____
Nc number of cold streams ____
Nh number of hot streams ____
P* vapour pressure ML" x T~ 2
Q heat supplied or rejected per unit
time ML2 T- 3
Qa Heat content of branch a ML2 T~ 3
VI
Qc cold utilities ML2 T~ 3
Qcmin minimum cold utilities ML2 T~ 3
Qc heat rejected at condenser per unit
time ML2 T~ 3
QH hot utilities ML2 T~ 3
Qnain minimum hot utilities ML2 T" 3
Qini Qout heat supplied and rejected from
a heat engine or heat pump ML2 T" 3
QB the total heat generated by the
reaction taking place at 25 °C ML2 T- 3
Qr heat supplied to reboiler ML2 T~ 3
Qxp the heat transferred across the
pinch ML2 T~ 3 Q- 1
q heat to vaporize one mole of feed
divided by molar latent heat ____
R reflux ratio ____
Rn minimum reflux ratio ____
S number of plates ____
Sn minimum number of plates ____
T absolute temperature 6
^Ta Maximum temperature that branch a
can reach without violating the
minimum approach temperature - the
initial temperature of branch a 6
Tact actual temperature 6
Tc condenser temperature 6
Ti temperature of interval i 6
VII
Ti+i temperature of interval i+1 6
Tint interval temperature 6
T* in minimum approach temperature 6
Tr reboiler temperature 6
Ts supply temperature 6
Tt target temperature 0
U»in minimum number of units including
heaters and coolers ____
V moles or mass of vapour stream in
the top of distillation column per
unit time MT~ 1
Xdi distillate composition of
component i ____
Xfi feed composition of component i ____
(XLK/XHK)<I the fraction of the light key
divided by the fraction of the heavy
key in the distillate ____
Greek Letters
cuv average relative volatility ____
as relative volatility at the bottom
of distillation column ____
ar relative volatility at the feed
stream ____
O.LK relative volatility of the light
key with respect to the heavy key ____
Vlll
or relative volatility at the top of
distillation column
6 the root of Underwoods equation
IX
CHAPTER ONE
Introduction
The topic of this research project is "The Integrated
Design of Chemical Plants With Energy Conservation", and
being so the work of the project is concerned with energy
integration in process design. The process of styrene
manufacture has served as a model process.
The main reason for choosing this process as the model
process is that styrene manufacture represents a typical
chemical manufacturing plant containing all the main aspects
of chemical plants. The process of styrene manufacture can be
divided into two sub-processes and these are;
1- Alkylation process for the production of ethylbenzene from
the reaction of benzene and ethylene.
2- Dehydrogenation process for the production of styrene from
the dehydrogenation of ethylbenzene.
Each of these processes has a reactor, distillation train
and necessary heat exchanger networks. The description and
the flowsheets of both processes are given in chapter 3 (mass
and energy balance).
1.1 Chemical Plants and Energy Requirements
To understand the energy requirements of chemical plants
it is helpful to examine the different components of a
chemical plant. A typical chemical plant can be represented
by the following diagram;
Raw _ materials
Separation process
We
Pure feed
1iste Un
Chemical reactor
* Separation process
reacted feed
Separation process
By product
Product
Figure 1.1 Typical chemical process.
As it can be seen from this diagram, a typical chemical
process is virtually the process of separation. For carrying
out the work of many separations energy must be provided.
Thus, the energy needjof chemical plants are usually high.
Traditionally the cost of energy was low in relation to
the selling price of final products and in relation to other
controllable costs. From 1973, energy costs soared and this
sudden increase in energy costs motivated industry to look
for energy savings. Many techniques have evolved for saving
and conserving energy in chemical processes and plants. For
existing plants energy savings can be achieved only by
revamping these plants. The revamp process is rather crude
and expensive. Therefore, energy saving techniques must be
incorporated into chemical plants at design stages.
However, no matter how advanced the unit operations to be
used in a chemical process are, they usually produce a poor
overall energy integrated process if these units are linked
up incorrectly. Therefore, the work carried out in this
research project is not into advanced unit operations, but
into the use of energy saving techniques in a systematic
procedure whereby a more elegant, sophisticated and overall
process energy integrated flowsheets can be generated.
1.2 Thesis Layout
The research work carried out in this thesis is divided
into six chapters excluding this chapter, and the text in
these chapters can be summarized as follows;
Chapter Two is devoted to the study of the principles of
process design incorporated with energy saving techniques,
and the theories that lie behind the work involved in the
following chapters. Also this chapter provides a review of
the work done in the area of process design and energy
integration.
Chapter Three provides the results of the mass and energy
balances that are carried out on the model process (styrene
process). These results define the original use of energy in
the process in terms of how much is consumed and how much is
released, so that process energy integration can take place
by meeting the energy released by the process energy
requirements. A computer program (physical properties data
service) has been used to estimate the properties needed in
these calculations.
Chapter Four deals with the integration of the distillation
sequence. Because many distillation sequences can be
generated for a multicomponent system**> (depending on the
number of the components fed to a separation process), and
because the proper integration of the heat exchanger network
of a distillation sequence has to be taken in the context of
overall process integration(2 > , then the optimum unintegrated
sequence that consumes less energy than the other sequences
is identified out of all the possible sequences. The
identified sequence is then energy integrated in the context
of overall process energy integration. Therefore this chapter
represents the first step towards the process integration,
since it is mainly concerned with the selection of the
optimum unintegrated sequence.
Chapter Five is concerned with heat recovery within the heat
exchanger network, and the use of different energy saving
techniques in order to maximize the recovery of heat. Thus,
the units interconnections are changed in order to produce a
network design that is compatible with the minimum use of
utility. However, the resultant network design is further
evolved so that more recovery of heat can be maintained, for
this pinch technology is adopted, and a computer software
(Target II) is used as a part of this work.
Chapter Six studies the interaction between the utility and
process design, the effect the utility has on the process
design and the external import of utility when the utility
introduced in different form. This has improved the heat
exchanger network efficiency, and thus a further reduction in
the amount of utility consumed is gained. Therefore, this
chapter represents a further step of energy integration.
Chapter Seven is mainly concerned with the effect of the
reduction of the utility consumed in the styrene process as a
whole on the steam and power system involved in the styrene
plant. Therefore, this system is examined before and after
energy integration is carried out on the process in terms of
fuel consumption, power generation and some relevant
characteristics.
Chapter Eight Summarizes the conclusions gained out of this
research work, also this chapter outlines some suggested
future investigations.
CHAPTER TWO
On the Process Design and Energy Integration
The "onion diagram" shown in Figure 2-1 represents the
hierarchical nature of chemical process design adapted from
Smith and Linnhoff< 3 > .
The traditional approach for process design is to tackle
the problem from the inner layer to the outer layer of the
onion step by step. Such an approach would build the inner
layers (reactor, separators) and seek a heat exchanger
network design to be fitted.
The present approach (pinch technology) takes the inner
layers (reactors and separators) as a start to calculate the
basic material and energy balance and define the process
streams. The results obtained from these balances will be
sufficient to design from the outer layer of the onion by
setting targets for minimum utility consumption. The next
step is to develop a design for a heat exchanger network that
is compatible with this minimum utility consumption. After
building the outer layers, the separators, heat engines and
heat pumps can be placed in the right position against the
process heat flow cascade. This will be explained in more
detail later.
In general energy saving techniques on a particular part
Figure 2.1 The hierarchical nature of chem'ical process design
in the process should not be considered in isolation of the
rest of the process, since good integration should take into
account each part in the context of an overall process.
Comprehensive literature reviews on separation process,
heat exchanger network and process synthesis with a view to
energy efficiency have been carried out by Nishida et al< 4 >,
Westerberge< 5 > and Gundersen and Naess< 6 >.
2.1 Process Synthesis
Process synthesis is a tool for building a flowsheet, and
is about choosing the best type and design of the process
unit and the best connection between process units* 7 ). The
connection between process units has been termed by Nishida
et al< 4) as process structure.
Two kinds of approach have been followed to synthesize a
process, the first one is beginning with a feasible flowsheet
and seeking to improve it, and the second one begins without
a flowsheet and starts synthesizing from scratch.
2.1.1 Synthesis With a Feasible Flowsheet In Hand.
This route of process synthesis is usually called an
evolutionary synthesis, since it deals with creating various
design modification to a previously synthesized process,
leading to an improved design* 8 > 9 >, This approach does not
guarantee a global optimum flowsheet, because this type of
approach can be thought of as a derivation of an improved
flowsheet from the original flowsheet by successive
modifications. Therefore, the better these initial flowsheets
are, the closer the final result will be to the optimum
solution. These initial flowsheets are generated by other
more general methods such as heuristic methods.
2,1.2 Synthesis Without having an Initial Flowsheet
Some advantage can be gained from synthesizing without
having an initial flowsheet such as, a global optimum can be
found because all flowsheets are possible and the nature of
the process can be understood very well. Synthesizing from
scratch can be based upon heuristic rules or a thermodynamic
approach.
2.1.2.1 Heuristic Approach.
Heuristics are a set of rules, usually known as rules of
thumb. This approach forms the oldest method of synthesizing
flowsheets when none exists, and they are based on practical
experience and the design of similar systems. Heuristics may
or may not give the best solution and provide no guarantee of
optimality because they are not based on a complete
understanding of the problem. Extensive lists of heuristic
rules for general process synthesis considerations, targets
and stream matching are given by Wells and Hodgkinson< 10 >.
2.1.2.2 Thermodynamic Approach
This approach is based on complete understanding of the
physics of the problem. Therefore this approach is rigorous
and usually give a very good answer* 1J 12 >. Pinch technology
is a good example of the application of a thermodynamic
approach, and is used through this research study.
2.2 Flowsheet Synthesis Problem
The flowsheet synthesis problem can be divided into four
areas of synthesis and these are; heat exchanger network
synthesis, separation system synthesis, heat and power system
synthesis, and reactor system synthesis. Thus the synthesis
of a complete flowsheet is found by choosing the best
designs, and connections between these four synthesis areas.
As has been mentioned earlier, process design starts with
10
the reactor, therefore fixing the reactor system by means of:
identifying the components in the raw materials, the purity
of the raw materials, the reaction conditions (temperature,
pressure) and kinetics, the catalyst, the optimum reactor
conversion...etc. would leave the problem of flowsheet
synthesis with three main areas.
2.2.1 Heat Exchanger Network Synthesis
Masso and Rudd< 13 > have stated that the design of heat
exchanger networks is a significant industrial design problem
aiming to reduce the overall utility consumption in a
processing plant. Heat exchanger network synthesis is a key
aspect of chemical process design* 12) . Such networks are
usually used to recycle thermal energy within a process
system preventing its wasteful loss with effluent materials.
This problem was first formalized by Masso and Rudd ( * 3 > , and
its goal was the development of a systematic procedure
capable of discovering the heat exchanger network which
reaches process specifications at minimum cost.
A typical chemical plant involves streams that have to be
heated (cold streams) or cooled (hot streams). The costs
involved include the cost of heating and cooling utilities
and the cost of the heat exchangers. The objective of
studying the design of heat exchanger networks is to heat and
11
cool the process streams from specified supply temperatures
to specified target temperatures at minimum total cost.
Development of systematic procedures to meet this
objective has been an active area of interest in the chemical
engineering for the last two decades. A review was given by
Nishida et al<«>.
Nishida et al< 4) included in his review the two major
developments and these are" network performance targets" and
"network temperature pinches". The location of the pinch was
described by Linnhoff et al< 14 - 15 > and Umeda et al< 16 . 17 >,
but the significance of the pinch in designing heat exchanger
network was not recognized in either source. The full
description of exploiting the pinch phenomena in heat
exchanger network design is given by Linnhoff and
Hindmarsh* x 8 > .
Targeting for minimum hot and cold utilities (maximum
energy recovery) and minimum number of units (heat exchangers
including heaters and coolers) is a very important task,
since the cheapest design is the one that has; firstly
maximum energy recovery, and secondly the minimum number of
units. Aiming for maximum energy recovery keeps the total
cost low because energy costs tend to dominate the total cost
of networks Sirola(19 >, and because energy costs tend to
dominate the total cost, the maximum energy recovery target
12
becomes the primary goal.
2.2.1.1 Heat Exchanger Network Specifications
The configuration of a heat exchanger network can be
presented as a set of process streams, and each stream has a
specified inlet and outlet temperature, flowrate and heat
capacity. The effect of temperature on the heat capacity is
usually negligible. Utility streams such as steam and cold
water are assumed to be available.
With consideration of minimum area for heat transfer, a
heat transfer coefficient should be specified.
All the information above can be taken from the process
flowsheet which is not yet heat integrated, as all the data
is based upon the mass and energy balances.
2.2.1.2 Different Approaches For Minimum Utility Usage
Minimum utility targets have been set in different ways.
One of the oldest approaches (an obvious way to set the
minimum utility target) is calculating the difference in heat
needed for heating the cold streams and the heat available
when cooling the hot streams. This approach has been
13
developed to account for cold streams which require heating
above any temperature available by the hot process streams or
any hot streams requiring cooling below any temperature
available by cold process streams* 20 • 21 >.
However, this approach has been refined to include the
minimum approach temperature, the stream levels and the
difference in heating and cooling duties* z2 23 >.
2.2.1.3 Heat Exchanger Network Representation
Throughout the development of heat exchanger networks
different representations have appeared. One of the oldest
representations is the "temperature-enthalpy diagram"* 24 >. In
this representation the temperature for each stream is
plotted against its enthalpy as shown in Figure 2-2. These
streams may be moved to the left or right because the
enthalpy values shown on the abscissa may be taken from any
datum point. Meat transfer between two streams is represented
by placing a cold stream (one which is to be heated in the
match) directly below a hot stream (one which is to be
cooled). The overlap, between the two streams represents the
heat transfer process. This heat flow is thermodynamically
feasible as the hot stream is hotter than the cold stream in
all places along the overlap. The vertical distance between
the streams is the temperature difference experienced along
14
the match.
The "simple match matrix" is shown in Figure 2-3. Pho and
Lapidus< 25 > introduced this representation which only allows
one match between any two streams.
The "heat content diagram" of Nishida et al< 26 > is the
third representation which is shown in Figure 2-4. In this
representation each stream is considered to be an area with
the vertical scale being the temperature and the horizontal
scale being the flowrate times the heat capacity of the
stream at that temperature. This area can be found by the
following equation;
Q = J mCp dT .......................................(2-1)
The area here represents the amount of heat to be
transferred. Thus an equal area in both one hot stream and
one cold stream is sectioned and numbered to represent the
match.
The fourth representation is given by Linnhoff and
Flower(22 >. As seen in Figure 2-5, in this representation the
process unit interconnections are directly represented. A
heat exchange process between two streams is represented by
placing a dumbbell on these two streams.
15
OJ
41ae
Steam
Hi H2 Heater
Enthalpy
Figure 2-2 Temperature enthalpy diagram
Hi H2 H3
Cl
C2
1
2
3
*r
Cl
Hi
Figure 2.3 Simple match matrix representation.
16
JUU
,o200
= 100o0.E ""200
150
mn
-12
•
Capacity flowrate
•/^/^
1 2
Heater
Figure 2.£ Heat content diagram
Hi
340 270
Fi'gure 2. 5 Temperature interval diagram17
2.2.1.4 Minimum Utility Targets and the Pinch Concept
Over the last ten years a procedure based on a strong
understanding of process thermodynamics has been developed to
guarantee minimum energy usage in the design of energy
integrated processes, the procedure is called "pinch
technology".
Pinch technology involves synthesizing an energy
integrated heat exchanger network design aiming for minimum
utility consumption. Targeting for minimum utility
consumption is the most convenient approach to get closer to
the optimum solution to the problem of synthesis of an
integrated heat exchanger network. Moreover, targeting gives
a designer confidence to attack his problem knowing what the
result should be, so the designer can modify his network
toward the target. The basic understanding of the pinch comes
from setting a target for minimum utility consumption and
this will be discussed thoroughly in the proceeding sections.
Minimum utilities may be calculated in two ways, via
graphical method, which was first proposed by Hohmann ( 27 > or
using the problem table method, which was proposed by
Linnhoff and Flower* 22 >.
18
2.2.1.4,1 Graphical Method
Any chemical process can be considered as a series of
streams requiring heating and cooling, This heating and
cooling is achieved by a combination of process interchangers
(matches between process streams) and heaters and coolers
supplied with utilities. A minimum approach temperature
( Tmin) is required for exchanging and recovering heat
energy. By combining all hot streams in terms of their heat
content (in temperature-enthalpy axes) the composite hot
profile will be obtained, and similarly by combining all cold
streams the composite cold profile will be obtained. The two
profiles thus portray the heat available and required for the
process.
Choosing the minimum approach temperature is mostly based
on experience, for instance, a minimum approach temperature
of 10 °C is found to be suitable and offer a good trade-off
between energy and capital cost in the case of liquid-liquid
heat transfer, while minimum approach temperature of 40 °C is
appropriate for gas to gas heat transfer< 28 > .
Now by choosing the minimum approach temperature it is
possible to bring the composite curves closest together and a
point of narrowest approach will occur. This point is called
the heat recovery "pinch" point.
19
To make it clear how the composite curves are constructed,
consider the two hot streams and two cold streams shown in
Table 2-1. These streams need to be changed from their supply
temperatures to their target temperatures consuming minimum
utilities (the data shown in Table 2-1 are taken from
reference 29).
A cold stream is a stream with its target temperature
higher than its supply temperature, for example, stream No.l
in Table 2-1 is a cold stream with a supply temperature of
20 "C and a target temperature of 135 °C. The heat capacity
flowrate is the product of the heat capacity times the
flowrate and it is given for each stream. The minimum
approach temperature is taken to be 10 "C.
Figure 2-6 and Figure 2-7 show how the hot composite curve
for the hot streams of Table 2-1 is constructed. Figure 2-6
shows the individual streams being plotted on a temperature
enthalpy axes, thus each arrow represents a hot stream. The
beginning and the end of each arrow represents the supply and
the target temperatures respectively. The vertical projection
of these arrows represents the streams heat duty, hence theto
slope of each arrow h minting the streams heat capacity flowrate.
The arrow can be shifted to the left or right without causing
any change in the calculations because there is no absolute
enthalpy, therefore the only thing that is important is the
vertical projection not its location.
20
Stream No. & Type
1 Cold
2 Hot
3 Cold
4 Hot
Supply Temp.rc)
20
170
80
150
Target Temp. CO
135
60
140
30
Heat Capacit Flowrate
2
3
4
1.5
Table 2-1 Process stream data
21
170°C
J50°C .<=>
4*
41a.E41
60° C
30° C
Enthalpy (KW)
Figure 2.6 Individual hot stream.
170°C 50°C
4*a.
60°C
30°C
Enthalpy (KW|
Figure 2.7 Hot composite curve.
22
In order to obtain the hot composite curve (a single curve
representing the heat content of all the hot streams) shown
in Figure 2-7, it is simply required to sum the heat
available in each of the temperature intervals. For example,
between 170 °C and 150 °C only stream No. 2 exists, and the
heat available in this temperature interval is the heat
content of stream No.2, so the composite curve has the same
slope as stream No. 2. Between 150 °C and 60 "C two hot
streams exist, so the total heat available is the sum of the
heat content of the streams, and the slope of the composite
curve in this interval is the sum of the slopes of both hot
streams. In the temperature-interval 60 °C and 30 °C only
stream No.4 exists, so the heat available in this temperature
interval is the heat content of stream No.4, and the
composite curve in this interval has the same slope as that
of stream No.4.
A cold composite curve can be constructed from the cold
streams in the same way as has been described for the hot
composite curve.
Hot and cold utilities can be found from Figure 2-8, which
contains both hot and cold composite curves. The open "jaw"
above the pinch determine the minimum hot utilities, and the
open "jaw" below the pinch determines the minimum cold
utilities. Figure 2-8 shows that the pinch divides the
profiles into two regions; a heat sink region which is above
23
o-o
2 a)a. E01
Minimum hot utilitynProcess below the pinch
Process above the pinch
Minimum cold utility
Enthalpy (KW!
Figure 2.8 Composite curves for the calculation of minimum utility consumption.
24
the pinch, and a heat source region which is below the pinch.
The area above the pinch represents an area of net def/cit of
heat requiring external hot utilities. The area below the
pinch has a net surplus of heat and thus needs external cold
utilities.
To guarantee that minimum hot and cold utilities have been
used, heat should not be transferred across the pinch.
Transferring heat from the sink area to the source area
requires extra heat in the sink area to cover the amount of
heat which has been transferred. In the same way the source
area will need extra cooling by the same amount of heat being
received across the pinch. The significance of the pinch is
summarized in three pinch design rules and these are;
1- Heat must not be transferred across the pinch.
2- Cooling must not be carried out above the pinch.
3- Heating must not be carried out below the pinch.
2.2.1.4.2 Problem Table Method
The problem table method is equivalent to the graphical
method. This method involves no trial and error and is based
entirely on simple arithmetic.
Data from Table 2-1 will be used once again to generate
25
the minimum utility targets and find the pinch location by
using the problem table method. This method can be divided
into three steps* 30 >;
1- The conversion of the actual stream temperatures to
interval temperatures:
For hot streams: Tint = Tact - ATnin/2 ...,.........(2-2)
For cold streams: Tint = Tact + &T«in/2 ............(2-3)
The resulted temperature intervals will be ranked and
duplicated points removed. Table 2-2 illustrates this step.
2- After setting the above temperature intervals, a heat
balance is carried out for the streams which fall within
each interval. This is done by using the following
equation;
A Hi = (Ti - Ti+l) (SCpc - SCph) ...................(2-4)
Now each interval has either a deficit or surplus of heat,
the results are shown in Table 2-3.
3- Heat should be transferred from higher to lower
temperature level, therefore the heat deficit of any
temperature interval can be satisfied either directly by
hot utilities or by cascading the heat surplus from a
26
Stream No. & Type
1 Cold
2 Hot
3 Cold
4 Hot
Ts —— »Tt (°C)
20 ——— »135
170 ——— >60
80 ——— »140
150 ——— »30
Temperature Interval
25
140
165
55
85
145
145
25
Setting Up Intervals
* ? t A 0 0 t * *
• • •J.40* * •
. . .140. . .
•••oO*«*
•••DD«»*
•••^0*»*
Table 2-2 Setting temperature intervals for the streams data of Table 2-1.
27
cold &Hot Streams
1
i i
i
i p
»
Interval Number
1 C C---loo- — -
1
—— 145 ———
2
—— 140 ———
3
O C-- — 85- — -
4
- 15 -
5
o e—-—-25—-—-
Ti-Ti + i
20
5
55
30
30
ZCpc — SCph
-30
-0.5
1.5
-2.5
0.5
Heat Surplus or Deff'cit
-60
-2.5
82.5
-75
15
Table 2-3 Heat balance for each interval
28
temperature interval at a higher temperature level.
By applying the cascading principle to the data shown in
Table 2-3, with the assumption that no hot utility is
supplied, the heat cascade shown in Figure 2-9a is obtained.
Each box in the heat cascade represents a temperature
interval, and the value written in the box represents the
heat surplus (indicated by a negative sign) or heat deficit
(indicated by a plus sign). However Figure 2-9a is still
infeasible because there is a negative heat flow which
opposes the Second Law of Thermodynamics. To overcome this
problem, the value of the largest negative heat flow should
be added to the hotest temperature interval in order to
obtain a non negative heat flow, as shown in Figure 2-9b.
The minimum hot utility is thus the smallest amount of
heat that must be put into the heat cascade to make all heat
flows nonnegative. Having done this a zero heat flow
represents the pinch location.
29
Temperature Interval(*C)
165
145
140
85
55
25
From Hot Utility (KW)
0
-60
60
-2.5
62.5
82.5
-20
-75
55
15
40
To Cold Utility
Fig.2-9a Infeasible heat flow cascade
From Hot UtilityW) 20
-60
80
-2.5
82.5
82.5
-75
75
15
60
To Cold Utility
Fig.2-9b Feasible heat flow cascade
30
Temperature Interval (°C)
165
145
140
85
55
25
From Hot Utility (KW)
1010
-60
70
-2.5
72.5
82.5
-75
75
15
60
To Cold Utility
From Hot Utility (KW)
20
1-60
1
80
r
-2.5
i
82.5P
82.5
i
0t
-75
1
35r
15
120
r
— » ——
1
40
To Cold Utility
Fig.2-10a Appropriate placement of heat above the pinch
Fig.2-10b Appropriate placement of cold utility below the
pinch
31
2.2.1.6 Grand Composite Curve
Figure 2-11 shows the derivation of the grand composite
curve from the heat flow cascade. Here the hot and cold
streams are combined and the heat surplus or deficit for each
interval is plotted on a temperature-enthalpy diagram* 31 >.
Hot and cold utility demands are represented by the width of
the open end at the top and bottom respectively.
The grand composite curve is very suitable for utility
selection* 32 > , it reveals that utilities do not have to enter
the process at extreme levels as assumed by the composite
curves. Thus, the grand composite curve can be well used as a
guide when a process needs utilities at different levels.
32
_.._--u c;----*---r-
55
25-——- ——— - ——
180
150
u ^1205**obo£ o»
60
30
Minimum hot utility
Process to process heat exchange
Process to process heat exchange
Minimum cold utility
10 20 30 40 50 60 70 80 90Enthalpy (KW)
Figure 2.11 The derivation of grand composite curve.
33
2.2.1.7 Constraints
In the pinch design method there are two constraints
affecting the way that stream matches are placed* 1 8 • 33 >. One
of them is related to the heat capacity flowrate for
individual streams, and the other is related to the number of
hot and cold streams and in both regions above and below the
pinch.
2.2.1.7.1 Heat Capacity Flowrate Constraint
At a pinch point the approach temperature is a minimum
for a given minimum temperature driving force. Away from the
pinch, the driving force can not be decreased i.e. the
temperature difference for heat transfer must be increased.
Therefore a match (between hot and cold streams) just
above the pinch requires a cold stream with heat capacity
flowrate higher than or equal to the heat capacity flowrate
of the hot stream, and this has been depicted in
Figure 2-12a. In terms of heat capacities the constraint can
be expressed by;
Cpc > Cph ...........................................(2-5)
If the hot stream heat capacity flowrate is greater than
34
Tl
IQ' c Ni
NJ
CT
~s a
2T
-*•
-1O
0)
D
O
Q°
3
u
n
Q-
"•"
-••
O
o n
-i ~^
(fl Q
n» •g D'
o IT
m 3 o »<
Tem
pera
ture
(°C
)i£
> C ro fo
Q
*&
?:r
(T>
O-i
IT
~*
SS.=
TQ
ro
:ii
a§.&
^s
o 0)
Q 3
O
Tem
pera
ture
(°C
m
that of the cold stream then the temperature difference
decreases away from the pinch and this violates the minimum
approach temperature Figure 2-12b.
The same principle must be applied to the region below the
pinch but here with the inequality reversed.
CP h > CP c .......................................... .(2-6)
2.2.1.7.2 Number of Process Streams Constraint
The pinch point divides the problem of heat integration
into two parts, one above the pinch and one below it. In the
part above the pinch the number of cold streams must be more
than or equal to the number of hot streams. This is so
because each hot stream needs to be cooled down to its pinch
temperature at the point of pinch. Therefore if there are not
enough cold streams to match the hot streams, heat will be
carried across the pinch and this would violate the minimum
utility target. This situation can be expressed by;
No > Nn .............................................(2-7)
The same principle can be applied to the part below the
pinch which requires that the number of hot streams must be
greater than or equal to the number of cold streams, thus;
36
Nh > Nc .............................................(2-8)
Finally the population of hot and cold streams has to be
such that it will allow an arrangement of exchangers
compatible with minimum utility usage.
If any of the above two constraints have not been
satisfied then the stream splitting technique must be
introduced.
2.2.1.8 Stream Splitting Technique
Stream splitting* is introduced when process stream data
at the pinch is not compatible with the design for maximum
energy recovery. Above the pinch the incompatibility happens
due to one or both of the following reasons;
1- Heat capacity flowrate for hot streams greater than the
heat capacity flowrate for cold streams.
2- The number of hot streams is greater than the number of
cold streams.
The hot stream that has a heat capacity flowrate greater
than any of the cold streams must be split into substreams
with smaller heat capacity flowrates. If the number of hot
streams is greater than the number of cold streams, a cold
37
stream must be split to produce more cold streams.
2.2.1.9 Minimum Number of Units
The design for minimum utility target should be continued
to achieve the minimum capital cost i.e. the minimum number
of units required* 34 • 3S > . The tick-off heuristic can be well
used to guarantee that the minimum number of units will be
used< 36 > .
The tick-off heuristic aim is simply to maximize the load
on an exchanger (containing two streams to be heat exchanged)
until one or both of the two streams is satisfied. Any
stream so satisfied is ticked-off and need no longer be
considered as a part of the remaining design task. This
heuristic resulted originally from Hohmanns equation for
targeting for the minimum number of units. The equation is of
the form;
U = N - 1 ...........................................(2-9)
where U is the minimum number of units (including heaters and
coolers) and N is the total number of streams including
utilities.
38
2.2.1.10 Utility Pinches
Utility selection often creates extra pinches, these
pinches are called utility pinches to distinguish them from
the original process pinch< 2 > . Utility heating above the
process pinch does not have to be placed at an extreme level.
It can be placed at lower levels anywhere above the process
pinch as long as it satisfies the target and keeps the heat
flow cascade positive. By increasing the use of utility at
lower levels, the heat flows could possibly be reduced to
zero thereby introducing a utility pinch.
Figure 2-13a shows the use of utilities at a temperature
higher than the highest temperature within the process, and
in this case high pressure steam is used. However not all the
process needs this utility at such a high level, so it is
more economical to make use of more moderate temperature
levels. For example, using steam at low pressure as shown in
Figure 2-13b, or if the high pressure was available in the
process and the process needs shaft work, then the available
steam can be used in turbine to produce sh&ft work and the
exhaust steam can then be used for process heating.
2.2.1.11 Threshold Problem
A threshold problem* 2 > occurs if the process does not have
39
T| ID' c n IT
1) o c 3 o a.
n> o
~J ( ! 1 1 1 li~
"
o ^
'L§
i&3
IL i i i 1
7> 1 1 1 1 1 1 r i i K 1 1
f> X
en (D n
Cfl 1 1 1 1 ± 1 1 h i 1
> X
r- O r»
«•
«
(S, I co
O *) ii O
co
O0
X O ro n X
N
1 5
o X D .A 1 <
k
0 X 1 1 X
Tem
per a
ture
(°C
)
T]
C ro
o>en
t*»N
l
c
0> o
1 1 1 10 4
— I
i. !
3
f 1 | 1
> n
, 1 1 1 1 1 I1 1 1
o <? D
n1 t
t> X
D •*.
< 0
C
> X
D
j i )
1
r> •5- o N>
1 1 1
^x
-L-x
°-
1 3
1 ——
M
5'S
i 1 1 1
X 3 3 0) ro"
< 0)m a-
o^ •o
Tem
pera
ture
(°C
)
O 3
a pinch. Such a process needs only one type of utility either
a hot or cold one. It is rare to find a threshold problem in
real life design because most industrial processes use
different types of hot and cold utilities.
Some threshold problems can have a large minimum approach
temperature before any increment in minimum approach
temperature requires any significant increase in the use of
utilities. The threshold problem is portrayed in Figure 2-14.
2.2.1.12 Process Improvement and Modification
Modification to a process can take two directions; either
by increasing the driving force between hot and cold streams,
or by decreasing the minimum hot and cold utilities.
Increasing driving force leads to reduction in capital cost,
since the increased driving force requires less area for heat
transfer.
The plus-minus diagram* 37 • 3S > is a general principle for
process modification, this is shown Figure 2-15. This diagram
shows that the directions for process modification can only
be achieved if the enthalpy changes of:
1- Hot streams above the pinch are increased.
2- Hot streams below the pinch are decreased.
3- Cold streams above the pinch are decreased.
41
Q.
e41
Steam
Enthalpy
Figure 2. U Threshold problem
42
4- Cold streams below the pinch are increased.
But often process temperatures are easier to change rather
than heat duties. Figure 2-15 shows that if the temperature
changes are kept on either side of the pinch, then the
driving force will be improved without effecting the energy
target.
If temperature changes go across the pinch then streams
will be shifted from one side of the pinch to the other, and
this will have a great impact upon utility target reduction.
The two principles for stream shifting across the pinch are;
1-Shift hot streams from below to above the pinch.
2- Shift cold streams from above to below the pinch.
2.2.2 Separation System Synthesis
A separation system involves a number of separation
processes, such as distillation, extraction,
absorption...etc. As the desired products in the styrene
plant (modelled in this study) have been purified via a
series of distillation columns, and as the distillation
process is the most widely used separation process in the
chemical process industries, synthesis of distillation
sequences will be looked at in more detail.
43
Oa
a.E
Shift hot stream
,,' ,Q Hm?n;'- -' \/
x
Shift cold stream
Enthalpy
Figure 2. 15 Plus minus principle on general composite curves
44
The synthesis problem for a distillation sequence has been
recognized over the last four decades and has been developed
since. Lockhart(39 > holds the first attempt to synthesize an
optimum distillation sequence. Lockhart found that the
removal of component should be done one-by-one as over head
product, and the most plentiful component should be removed
first.
The energy consumption in the distillation sequence was
the main concern for Harbert(40) , so he recommended that the
most difficult separation should be saved for last, and also
found that a 50-50 split is quite economical. Total vapour
flow has been considered by Rod and Marek ( 4 1 > as a factor
that has great effect on economic criteria, so they advised
to sequence with minimum total vapour flow. Heaven* 42 > has
studied the synthesis problem of distillation sequence in
more detail and has generated heuristics similar to the ones
above, except he pointed out that high recovery fractions
should be done last.
Some observations have been made by Freshwater and
Henry< 43) when they undertook an extensive study on the
applications of the heuristics suggested by Heaven. Their
observations are; it is favourable to remove components in
decreasing order of volatility as over head product if there
is no difficult separation, and to leave the most difficult
separation last was found in several cases to be desirable.
45
Energy saving distillation configurations such as, a
single column with a side-stream product, and a
prefractionator column followed by a side stream column, as
well as the traditional configuration column have been
studied by Doukas and Luybeen* 44 '. They pointed out that when
the least volatile component is less than 10% then a single
side-stream column is the most economical. They also found
that utility cost is the most dominant factor.
There are six heuristics that have been suggested by
Seader and Westerberg* 4 5 > , to help in building up an initial
sequence and give a basis for later improvement. These
heuristics have been given by different authors before, but
Seader and Westerberg have arranged these heuristics in a
procedure to be followed one by one, they are as follows;
1-When the relative volatility between the key components is
less than 1.05, the split has to be indentified and is
considered to be forbidden.
2- Easiest separations (big difference between the components
relative volatilities) should be done first.
3- When the mole percentage of the feed components varies
widely but the relative volatilities do not, the removal
of the components must be done in the order of decreasing
molar percentage in the feed.
46
4- The direct sequence should be followed when neither
relative volatility nor molar percentage in the feed
varies widely.
5- When a mass separation agent is used (e.g. extractive
distillation), the agent should be removed in a separator
immediately following the one into which it is introduced.
6- When multicomponent products are specified, a sequence
of separation should be followed that produces the
smallest product set.
Heuristics 1 and 6 have been used by Thompson and
King(46 > 47 >; 2 and 3 have been used by Heaven* 42 >; 4 and 5
have been used by Freshwater and Hendry* 4 3 > and Hendry and
Hughes* 48) respectively.
A similar approach has been presented by Nath and
Motard(49) . One of the important features in their heuristics
is that some of them provide a guideline for the design of
the separators of the initial sequence such as; operating
pressure should be close to ambient, and operating reflux
equal to 1.3 times the minimum should be used for each
column.
47
2.2.2.1 Energy Integrated Distillation Column
The distillation process is a very highly energy-consuming
unit operation. With some processes consuming a third or more
of their energy in distillation alone. For this reason energy
integration and optimization have been carried out by many
workers* 50 • 53 >. Energy integration can be within the
separation system or with other parts of the process.
Different techniques have been applied for better energy
efficiency, for example thermal coupling* 54 > , multiple effect
column* 55 >, vapour recompression* 56 >, intermediate condenser
or reboiler< 57 > , side-stream stripper or rectifier< 58 >, and
many others. All the above traditional techniques have taken
the individual column in isolation independently from the
rest of the process, and integration without a complete
understanding of the background of the process.
Linnhoff et al< 2 > , and Smith and Linnhoff< 3 > have
presented a technique that takes the integration of an
individual column into context with the heat integration for
the overall process. They have concluded that, good
integration can result in a column operating at effectively
zero utility cost, although some of the traditional schemes
seem good on stand alone basis they can in certain
circumstances be counter-productive when it comes to column
integration with the rest of the process, and for energy
48
integration there is no need for using complex column
arrangements as simple columns can provide good integration.
In a case where heat flows are limiting integration
possibilities or the distillation system can not be
integrated with the rest of the process for operability or
other constraints, then complex column arrangements should be
considered.
This approach which takes the integration of each
distillation column in the context of the overall process is
based mainly on the principle of proper placement which
depends heavily on the position of the distillation column
relative to the heat recovery pinch.
2.2.2.2 Appropriate Placement of Distillation Column
To run a conventional distillation column, heat (Qr) has
to be supplied in the reboiler at temperature (Tr), and heat
(Qc ) has to be rejected in the condenser at a lower
temperature (Tc). The temperature levels of these heat loads
(Qr and Qc ) assigns the position of the distillation column
relative to the heat recovery pinch.
There are two possible placements for distillation columns
within the process heat flow cascade, either the condenser
49
and the reboiler temperature span the pinch temperature or
they do not. If the separator is across the pinch then the
heat Qr is received at a temperature higher than the pinch
temperature, and the heat Qc is released at a temperature
below the pinch temperature as shown in Figure 2-16a. This
means that heat has been taken from the part of the process
which is a heat "sink" and given to the part of the process
which is a heat "source" resulting in an extra hot and cold
utilities. Integrating this separator in this manner can
never give an advantage over the stand alone basis.
If the column is not across the pinch, then the column
could be either above or below the pinch as depicted in
Figure 2-16b. When the column is entirely above the pinch,
the heat Qr will be taken from the part of the process which
is above the pinch (sink) and given back again as Qc at lower
temperature above the pinch. Therefore the utility in this
case would not change and the column is run for free (no
utility consumption). Below the pinch the situation is
similar.
For operability reasons, the reboiler and the condenser do
not both have to be integrated with the process, because this
can make the column difficult to start-up and control. Above
the pinch the reboiler can receive its heat directly from the
hot utility and the condenser should be integrated with the
rest of the process see Figure 2-17. Below the pinch the
50
QHaln + Qr
Pinch
QH.in + (Qr - Qc )
Qc • 1 n +Qc
Fig.2-16a Distillation column Fig.2-16b Distillation column across the pinch not across the pinch
51
QltMin - QC
Pinch
Qc
Fig.2-17 Integration of distillation column with consideration of operability and control.
52
reboiler should be integrated with the rest of the process
while the column can reject its heat directly to the cold
utility see Figure 2-17.
2.2.3 Heat and Power System Synthesis
A power plant, working continuously on full load may have
an overall efficiency of 30%, this is low due to the large
percentage of energy which is rejected to the condenser* 59 > .
The power plant efficiency can be considerably increased if
it is part of a plant which requires a low temperature supply
of heat for heating or process work. The steam power system
in this case must generate steam at certain pressure and
temperature levels.
Due to the different requirements imposed by the process
systems depending on the various operation modes, the steam
power system has to have a greater requirement for
flexibility and reliability* 6 °> .
Nisho and co-workers* 6 1 > called the process which needs a
large amount of heat relative to power is "steam dominant
case", and if opposite situation was the case then the
process called "power dominant case". Hence their proposed
design of steam power system was dependent on the processes
relative needs of heat and power.
53
2.2.3.1 The Integration of Heat and Power in the process
Network
This subject covers many areas such as site combined heat
and power, on plant power generation, heat pumps, and
refrigeration systems.
Earlier attempts to this subject were few and based on
experience, for example Menzies and Johnson* 63 > . A
significant new insight has been gained based on extensions
of the fundamental concept of heat exchanger network
synthesis and the process heat flow cascade* 62 • 64 • 65 >. This
new insight depends heavily on the position of the heat
engine and the heat pump relative to the pinch.
2.2.3.1.1 Appropriate Placement of Heat Engines
A heat engine can be thought as a distillation column, it
takes heat at higher temperature and rejects heat at lower
temperature except that the heat engine produces work.
Therefore, the appropriate placement of heat engine can not
be across the pinch, because heat would be transferred from
above to below the pinch causing extra hot and cold
utilities.
54
Integrating such an engine across the pinch can never be
better than the stand alone situation (Figure 2-18a).
Placement above or below the pinch would give an efficiency
of 100% (Figure 2-18b).
2.2.3.1.2 Appropriate Placement of Heat Pumps
The principle of appropriate placement can also be applied
to heat pumps. Heat pumps consume heat at a low temperature,
the stream supplying this heat undergoes work, and is
rejected at a higher temperature. The heat rejected is
equivalent to the heat input and work done. Now placing a
heat pump above the pinch will not give an energy saving but
conversion of work into heat at higher temperature. Placing a
heat pump below the pinch would result in the conversion of
work into heat at higher temperature intervals and this will
be degraded to the cold utilities, and thus increasing the
cold utilities by the amount of work done (Figure 2-19a).
Therefore, energy saving can not be obtained unless the heat
pump is placed across the pinch.
Having placed the heat pump across the pinch the heat
would be taken from below the pinch causing a reduction in
cold utilities, and rejected with the work input above the
pinch giving a reduction in hot utilities (Figure 2-19b). An
efficiency of 100% can be gained by placing the heat pump
55
Pinch
QH»in Qin1
1Qi
1 ,Q2
,0
1Q4
r
1
Qin-
\ f
Heat Engine
———
\ I
Qs +Qi n -W
|Qc«in+Qin-W
Heat Engine
(Q-W)=QH.ln
Heat Engine
Qcain
Qc»in-W
Fig.2-18a Heat engine across Fig.2-18b Heat engine not the pinch across the pinch
56
QH»in-W
Pinch
QH«in-(Qin+W)
w
w
,+w tQcBin-Qin
Fig.2-19a Heat pumps not across the pinch
Fig.2-19b Heat pumps across the pinch
57
across the pinch.
2.2.3.2 Selection of The Right Heat Engine
Townsend and Linnhoff< 65 > have presented a method for the
selection of the right heat engine depending on the process
profile. If the exhaust of a heat engine is used for process
heating then the type of heat that would be received by the
process depends on the type of the working fluid in the heat
engine. By using a gas turbine the heat would be sensible
heat and this can be represented on a temperature-enthalpy
diagram as sloping straight line. By using a steam turbine
the heat would be latent heat and can be represented on a
temperature-enthalpy diagram as a horizontal line. If the
process profile above the pinch was as in Figure 2-20a
(a nearly straight slopping line), then the gas turbine would
be a good choice. This is because the loss of minimum driving
force will small and maximum work output will be gained.
Whereas, the steam turbine is a good choice for the process
profile shown in Figure 2-20b. To reduce the loss of driving
force and maximize work, a multistage turbine that can
produce a number of pressure levels can be used to match the
process profile.
58
34J03In 0)
Enthalpy
Figure 2-20a Chosing the gas turbine to match the process,
34->US h 0)
Steam
Steam
Steam
Enthalpy
Figure 2-20b Chosing the steam turbine to match the process,
59
CHAPTER THREE
Mass and Energy Balance
The work in this chapter involves mass and energy balances
around both the entire plant and the main parts of the plant.
Information on process parameters such as, flowrates and
concentration, can be obtained from the calculations of the
mass balance. The amount of energy consumed and released in
the process can be obtained from the calculations of the
energy balance.
The data resulting from the mass and energy balances are
important since they represent the process stream data, which
are used in setting targets for minimum hot and cold
utilities, and synthesizing an optimum heat exchanger network
by integrating all available heat sources and sinks in the
process.
3.1 Mass Balance
The results from the mass balance are given in this
chapter. A sample of the calculations is shown in appendix A.
60
3.1.1 Overall Mass Balance
The plant has been designed to produce 500,000 tons of
styrene per year. The product purity to be maintained is as
follows;
Component Wt. per cent
Styrene 99.7
Ethylbenzene 0.3
The finished product is thus made up of the following
quantities:
Styrene = 498500 ton
Ethylbenzene = 1500 ton
Working days in a year are taken to be 300 days, and the
basis for the calculations of mass balance is taken to be 1
hour. On this basis the amount of finished products to be
produced are:
Styrene = 69236 kg per hour
Ethylbenzene = 208 kg per hour
Total = 69444 kg per hour
A schematic diagram of the overall mass balance is given
in Figure 3-1. As described previously the process of styrene
manufacture includes two main process steps, namely
alkylation of benzene with ethylene to give ethylbenzene and
61
Vent gas = 6245.5 kg
Bz = 118018.4 kg
Et = 25421.2 kg
Steam=485151.4 kg
Styrene Plant
Styrene = 69444.4 kg—————————————p.
Bz = 67600.5 kg
Toluene = 2058.4 kg
Water = 482535.6 kg Tar = 360 kg
Figure 3.1 Overall mass balance of the entire plant,
62
the dehydrogenation of ethylbenzene to give styrene.
3.1.2 Alkylation Process Step
When ethylene and benzene react in the presence of
aluminium chloride and hydrogen chloride, alkylation of the
benzene ring occurs to produce ethylbenzene and higher
ethylated benzenes. Aluminium chloride is used as an
alkylation catalyst. In order to operate at high catalyst
efficiencies, hydrogen chloride must be added as a
promoter* 66 ).
Alkylation takes place at a temperature and pressure of 95
°C and 1.3 bar respectively. With a molar ratio of ethylene
to benzene of 0.6, it is reported* 67 > that the following
product composition was maintained at optimum operating
conditions:
Component Wt. per cent
Benzene 40.6
EB 46.8
DEB 11.9
HPEB 0.72
where EB is ethylbenzene, DEB is diethylbenzene and HPEB is a
polyethylbenzene that is at a grade higher than
diethylbenzene such as triethylbenzene, tetraethylbenzene and
63
so on.
In a typical operation the product from a reactor is
cooled and passed into a settler where the hydrocarbon is
decanted from the catalyst, and the catalyst layer is
recycled to the reactor. The hydrocarbon is washed with water
and caustic to remove traces of catalyst. After this, crude
ethylbenzene is sent to a distillation train for recovery of
product ethylbenzene. The flowsheet of the alkylation process
is given in Figure 3-2.
3.1.2.1 Mass Balance Over the Alkylator
The results of the mass balance are shown in Figure 3-3.
Recycled PEB 20298.7 kg
Bz = 118018 kg
Et = 25421 kgAlkylator Crude EB = 163738 kg
Fig.3-3 Alkylator Mass Balance
As can be seen from the figure, the alkylator produces
163738 kg per hour of crude ethylbenzene and this is purified
64
Rec
ycle
d P
EB
O)
Ol
W.S
. W
ASH
ING
SYS
TEM
Et -
ETH
YLE
NE
BZ
- B
EN
ZEN
EEB
-
ETH
YLB
EN
ZEN
ED
EB
-DIE
THY
LBE
NZE
NE
PE
B - P
OLY
ETH
YLBE
NZE
NE
HP
EB
-HIG
HE
R P
OLY
ETH
YL B
ENZE
NE
Res
idue
EB
Fig
ure
3.2
Eth
ylbe
nzen
e pr
oces
s flo
wsh
eet.
in a series of distillation columns to give ethylbenzene of
desired specifications for the production of styrene through
the dehydrogenation of ethylbenzene.
3.1.2.2 Mass Balance Over the Distillation Process
The components that are recovered from the separation
process are benzene, ethylbenzene and polyethylbenzene.
Benzene and polyethylbenzene are recycled to the alkylator,
ethylbenzene is sent to the dehydrogenation process for
styrene production. Table 3-1 shows the results of the mass
balance around each column.
3.1.3 Dehydrogenation Process Step
In this process, ethylbenzene produced from the alkylation
process will be dehydrogenated to produce styrene and some
side-products. The main reaction of this process is
630°C CeHsCzHs ———————————+ CeHsCzHa + Hz
In addition to this, 15 other side reactions are possible,
and of them five are important* 68 > , and given by
CcHsCzHs ————————* CeHe + CzH4
CeHsCzHs + Hz ————————»• CeHsCHa + CH4
66
Column Name
Peed (kg)
Top
Product
(kg)
Bottom product
(kg)
Stri
ppin
g Co
lumn
Benzene
Column
Ethylben-
zene Column
Polyethyl
Still
Benzene
66477.7
Ethylbenzene
76629.5
Diethylbenzene
19484. B
Higher
1211.6
Polyethyl benzene
Benzene
66477.7
Ethylbenzene
76629.5
Diethylbenzene
19484.8
Ethylbenzene
76629.5
Diethylbenzene
19484.8
Higher
1211.6
Polyethylbenzene
Benzene
Ethylbenzene
Diethy
lbenze
ne
Benzene
Ethylbenzene
Diethy
lben
zene
Triethylbenzene
66477.7
7662
9.5
19484.8
6647
7.7
76629.5
385.
1
1198.9
High
er
Poly
ethy
lben
zene
1211
.6
Ethylbenzene
Diet
hylben
zene
Diet
hylben
zene
Resi
due
7662
9.5
19484.8
19099.7
12.7
Table
3-1
The
resu
lts
of the
mass
ba
lanc
e ov
er the
separation process
in the
alkylation st
ep.
1/2C2H4 + HzO —————————> CO + 2Hz
CH4 + H2O ————————+ CO + 3H2
CO + H20 ————————»> CO2 + H2
The dehydrogenation reaction is carried out at a
temperature of 630°C and a near atmospheric pressure. The
main reaction is the only reaction that is found to be
reversible and as the right hand side of the reaction has two
moles of product (styrene and hydrogen), a shift towards the
left is more favourable. To prevent this, steam is added to
the reactor. The effect of this, is to reduce the styrene
partial pressure and to increase the molar conversion of
ethylbenzene thereby favouring the forward reaction. Also
this steam is used to supply the heat that is required by the
endothermic dehydrogenation reaction.
A furnace is used to superheat the steam required.
Superheated steam in the molar ratio of 15 steam to 1
hydrocarbon is used with conversions of 35 to 40% per
pass* 6 6 ) .
The ethylbenzene entering the dehydrogenator must contain
less than 0.04% diethylbenzene, since this material is
partially converted to divinylbenzene which polymerizes very
rapidly to form insoluble residues in the purification
system.
68
The dehydrogenation process requires four distillation
columns to separate the reactor effluent. A flowsheet of the
process is shown in Figure 3-4.
3.1.3.1 Mass Balance Over the Dehydrogenation Reactor
The results of the mass balance are shown in Figure 3-5
Recycled EB
EB
EB = 190754 kg
Steam= 485151 kg
DehydrogenatorCrude styrene = 187124 kg Gases = 6245 kg Steam = 482535 kg
Fig.3-5 The dehydrogenator Mass Balance
3.1.3.2 Mass Balance Over the Distillation Process
The components that are recovered from the separation
process are benzene, toluene, ethylbenzene and styrene. The
benzene is recycled to be used in the alkylator. The toluene
is sold as an intermediate product for which there are
various uses e.g. for making explosives. The ethylbenzene is
recycled to the dehydrogenator, and styrene is the final
product. Table 3-2 shows the results of the mass balance over
the distillation sequence.
69
Jl UD"c —i0>
CJ
o0) IT ^Q.
3UD rt>D Q5'
o o
VI
o
Furnace
o o
Dehydrogenator g>
o en ^ <•* c n 0.0» 3
COn a3
en
cno Benzene Toluene column
1
•Jk•33n n«< o_na.mCD
tn Ethylbenzene column a>03M
•<_£5'
3
O O
8to
1 Styrene column
en
tt(9
m Benzene column
Column na
me
Benzene -Toluene
Column
Ethylbenzene
Column
Styrene
Column
Benzene
Column
Peed (kg)
Styrene
69236.1
Ethylbenzene
114333.1
Toluene
2058.3
Benzene
1122.7
Tar
209.5
Styrene
69236.1
Ethylbenzene
114333.1
Tar
209.5
Styrene
69236.1
Tar
209.5
Ethylbenzene
208.3
Benzene
1122.7
Toluene
2058.7
Top
Prod
uct
(kg)
Toluene
2058.3
Benzene
1122.7
Ethylbenzene
114124.6
Styrene
69236.1
Ethylbenzene
208.3
Benz
ene
1122.7
Toluene
3.8
Bottom Pr
oduc
t (k
g)
Styrene
69236.1
Ethylbenzene
114333.1
Tar
209.5
Styrene
69236.1
Tar
209.5
Ethylbenzene
208.3
Tar
209.5
Toluene
2054.6
Table
3-2
The
results
of the
mass balance
over the
separation process
in the
dehydrogenation process
3.2 Energy Balance
The results of the energy balance are presented in this
chapter. A sample of the calculations is given in appendix
B.
3.2.1 Alkylation Process Step
The reaction in the alkylator is exothermic, and the heat
generated amounts to '8 MW.
Crude ethylbenzene after coming out from the alkylator
passes through a heat exchanger and is cooled from 95 °C to
40°C. The heat removed in the exchanger is 4.7 MW, After
leaving this heat exchanger the stream goes through a
combined washing process and treatment step to remove the
catalyst aluminium chloride , and then the stream is sent to
a separation step (four distillation columns). Table 3-3
gives the results of the energy balance carried out for the
individual columns in the distillation train.
In this process step there are 3 coolers. The location of
each individual cooler and the amount of heat removed from
each one are as follows;
1- Cooler after benzene column to cool down benzene product
72
w
Column Name
Str ip
ping
Column
Benzene
Column
Ethylbenzene
Column
Polyethyl
Still
Feed Heat
Content
(MW)
1.211
25.179
6.55
0.115
Top
Product
Heat Content
(MW)
25.179
1.846
4.637
0.069
Bottom Product
Heat Content
(MW)
0.114
6.55
1.934
0.001
Heat Required
in the
Reboiler
(MW)
24.083
12.791
29.155
0.104
Heat Rejected
in the
Condense
r (MW) ——
29.57
29.137
0.104
Table
3-3
The
results
of the
energy balance
around each column
in distillation train
in alkylation
process.
to 40°C, and the amount of heat removed is 1.34 MW.
2- Cooler after ethylbenzene column to cool down ethylbenzene
product to 40 °C, and the amount of heat removed is 4 MW.
3-Cooler to cool down the recycle to 40*C, and the amount of
heat removed is 1.8 MW.
3.2.2 Dehydrogenatlon Process Step
The reaction in the dehydrogenator is endothermic, and the
heat required by the reaction is win* MW.
Two heat exchangers are used to cool down the effluent
from the dehydrogenation reactor. In the first heat
exchanger, the effluent is cooled from 565° to 441.5 °C and
the heat removed is 52.7 MW. In the second heat exchanger,
the effluent is cooled from 441.5° to 306 °C and the heat
removed is 54.9 MW. After the two heat exchangers the
effluent stream goes through a cooler and condenser, and then
a gravity separator before it goes to a distillation train.
At the cooler, the stream is cooled down to 105 °C losing
73.4 MW. At the condenser the stream loses a further 320 MW.
The gravity separator cools down the stream after
condensation to 49 "C and decants the hydrocarbons from the
water phase.
74
In the dehydrogenation step, there are two heaters. The
first heater is to heat up the feed to benzene-toluene
column, and the heat added is 2.43 MW. The second heater is
to heat the feed to the ethylbenzene column, by adding 1.34
MW of heat. Also there is a cooler after the ethylbenzene
column and the heat rejected in this cooler is 1.23 MW. Table
3-4 gives all the results of the energy balance carried out
over each column in the distillation train.
3.3 Concluding Remarks
The results of the energy balance are the main objective
of this chapter (for the reasons mentioned earlier in this
chapter). The overall mass flowing through the process is
considered to be constant, since no change will occur to it
as the process is energy integrated. Therefore, the mass
balance has been carried out in order to facilitate the
energy balance.
The hot and cold utilities supplied to the alkylation
process are about 66 and 71 MW respectively. The hot and cold
utilities supplied to the dehydrogenation process are about
221.6 and 500 MW respectively. In both processes a high
percentage of these utilities are consumed by the
distillation trains. In the alkylation process the reboilers
and condensers associated with the distillation train consume
75
Colu
mn Name
Benzene -toluene
Colu
mn
Ethylbenzene
Colu
mn
Styrene
Column
Benzene
column
Feed Heat
Content
(MW)
4.668
8.216
1.728
0.048
Top
Product
Heat Content
(MW)
0.048
1.837
1.074
0.03
Bottom Product
Heat Content
(MW)
6.879
2.96
0.005
0.098
Heat Required
in th
e reboiler
(MW)
6.89
82.987
15.226
0.583
Heat rejected
in the
condenser
(MW)
4.63
86.404
15.875
0.5
Table
3-4
The
results
of the
energy balance
around each column in distillation train
in
dehydrogenation process.
about 66 and 58.8 MW respectively. In the dehydrogenation
process the reboilers and condensers associated with the
distillation train consume about 105.7 and 107.4 MW
respectively.
Therefore, the energy integration pursued in the next
chapter will be carried out on the distillation trains in
both processes in order to select the separation sequence
that consumes less energy than the other possible sequences.
77
CHAPTER FOUR
The Selection of an Optimum Unintegrated Distillation
Sequence
The problem of synthesizing an integrated distillation
sequence can be decomposed into two steps. Step one is to
identify the best unintegrated sequence, since the optimal
unintegrated sequence tends to possess the largest potential
for heat integration* 69 > . Heuristic rules are found to be
compatible with the selection of the optimal (or near
optimal) unintegrated sequence* 69 ~ 7 1 > , thus heuristics are
adapted to distinguish the sequence (or few sequences) that
stands out amongst all the possible separation sequences.
The second step is the design of the optimum heat
exchanger network for the identified sequence. This step has
been detailed in the preceding chapters, since the
integration of the optimum heat exchanger network is taken in
the context of the integration of the process as a whole.
Therefore, this chapter is mainly concerned with the
selection of the optimum unintegrated distillation sequence.
4.1 Heuristics Used
Many heuristics have been generated, and these can be
78
summarized by the following four< 69 » 72 ):
1- Perform difficult separation last.
2- Largest component should be removed early in the sequence.
3- Favour near equimolar split between the top and bottom
products.
4- Favour the direct sequence (sequence which remove the
components one by one in column overheads in decreasing
order of volatility).
These heuristics have been considered under conditions
where conventional columns are in use, and no heat
integration is involved (i.e. all reboilers and condensers
are serviced by utilities).
4.2 Alkylation process
4.2.1 Identification of Possible Unintegrated Sequences
The possible number of sequences for separating
J-components in a system operating conventional distillation
columns is given by the following equation* l >;
Number of sequences = ————————— .....,...........(4-1)
The number of components that have to be separated in the
79
alkylation process is four, and these components are shown in
Table 4-1. Applying the equation above, the number of
possible sequences will be as follows;
Component mol% B.p.(C')
Benzene
EB
DEB
HPEB
49
42
8.5
0.5
80.1
136.3
183
over 220
Table 4-1 The composition of the feed to separation system in alkylation process.
(2(4-1))!Number of sequences = ——————————— = 5
4! (4-1)!
The sequences are shown in Figure 4-1. These sequences
differ in their energy demand, hence one of these sequences
may be selected on this basis. The selected sequence should
consume less energy than the other sequences, and require
relatively less capital. The selection procedure is effected
by the direct application of the heuristics mentioned, the
philosophy which lies behind these heuristics and finally by
calculating the energy consumption for each sequence.
To simplify the synthesis procedure some assumptions have
been made which are suitable for the process under study
80
BZ EB PEB
BZ.BB.DJB, i°' HPEB ^^
81°
c
"o o
148°
EB, DEB, HPEB
... BZ,EB,DEB ID) HPEB —
|BZ81°
c E"o o
148°
EB.DEB.HPEB
BZ.EB.DEB (C) .HPEB -*"
BZ117°
c E3 00
195°
EB
|138°OJc30 O
195°
|
1185*
C
°o o220°
DEB', HPEB
151°
c E
_3
220°
1 __
PEB
Residue
BZ81°
CNCe
"o O
U5°
J EB
DEB, HPEB BZ.EB.PEB BZ
... BZ.EB.DEB t a l HPEB
IP i BZ EB DEB19 ' HPEB -*•
132°
c
~o o
220°
Residue
BZ, EB.PEB132°
c
30 0
220°
81°
c £"o o
148°
EB, PEB
BZ117°
c E
_30 O
195°
EB
12
IS
If
2
Residuetheavy materials1 EB18°
c_3"o U5°
1PEB
1 PEJ5°
ce3 OU20°
B~
Residue
EB
138°
roCE30 O
195°
PEB
BZ81°
CO
cE3 Oa
U5°
Residue PEB EB
Figure 4.1 The possible distillation sequences for separating the effluent of the alkylator.
81
(some of these assumptions were also used by some other
investigators in their work< 7 * > 7 3 > . The assumptions are;
1-Simple columns are considered in all sequences (simple
column is one which separates a feed into two product
streams).
2- High recovery of the key components.
3- The operating pressure in each column is taken to be near
atmospheric. In the styrene process there is no specified
difficult separation that requires the column to run under
vacuum. Furthermore, a high pressure operation will result
in high temperatures on the condenser and reboiler and
consequently utilities at higher temperature levels are
needed. Therefore, atmospheric pressure is compatible with
both the materials being separated and energy conservation
principles.
4- Partial condenser will not be used, because the total heat
rejected will not reflect the right amount needed to
measure the differences between the sequences in terms of
cold utility consumption.
5- No heat integration is taken into consideration, since
heuristics can only be applied for columns that do not
involve heat integration.
82
6- Heating and cooling for the intermediate streams are
negligible compared with the total heating and cooling
loads.
4.2.2 Heuristics Application
The application of the heuristics mentioned in section 4.1
to the sequences in Figure 4-1, should generate one or a few
sequences that stand out among the other sequences.
It is very clear from Table 4-1 that there is no difficult
separation, since there are no close boiling components. This
would make no use of heuristic No.1. Heuristic No.2 indicates
that the sequence in Figure 4-la is the one that should be
chosen, since the largest components, which are benzene and
ethylbenzene, are removed earlier in the sequence. Applying
heuristic No. 3 would nominate the sequences in Figure 4-la
and b, because the equimolar split occurs at the first
column. Heuristic No. 4 would also select the sequence shown
in Figure 4-la, since the components in this sequence have
been removed in decreasing order of volatility. Table 4-2
summarizes the application of the heuristics to the sequences
shown in Figure 4-1.
The conclusion drawn from this, is that the sequence shown
in Figure 4-la can be considered as the best unintegrated
83
6- Heating and cooling for the intermediate streams are
negligible compared with the total heating and cooling
loads.
4.2.2 Heuristics Application
The application of the heuristics mentioned in section 4.1
to the sequences in Figure 4-1, should generate one or a few
sequences that stand out among the other sequences.
It is very clear from Table 4-1 that there is no difficult
separation, since there are no close boiling components. This
would make no use of heuristic No.1. Heuristic No.2 indicates
that the sequence in Figure 4-la is the one that should be
chosen, since the largest components, which are benzene and
ethylbenzene, are removed earlier in the sequence. Applying
heuristic No. 3 would nominate the sequences in Figure 4-la
and b, because the equimolar split occurs at the first
column. Heuristic No. 4 would also select the sequence shown
in Figure 4-la, since the components in this sequence have
been removed in decreasing order of volatility. Table 4-2
summarizes the application of the heuristics to the sequences
shown in Figure 4-1.
The conclusion drawn from this, is that the sequence shown
in Figure 4-la can be considered as the best unintegrated
83
Heuristic(1)
Heuristic (2)
Heuristic (3)
Heuristic (4)
Sequence (a)
Yes
Yes
Yes
Sequence (b)
No
Yes
No
Sequence (c)
No
No
No
Sequence (d)
No
No
No
Sequence (e)
No
No
No
Table 4-2 The application of the heuristics to the sequences shown in Figure 4-1.
84
sequence, as far as the heuristics are concerned.
4,2.3 Heuristics Philosophy
The four heuristics relate directly to the flowrate of
components in distillation columns. They usually minimize the
flowrate of the components. The minimization of the flowrate
keeps not only the energy cost down but also the capital
cost. It keeps the energy cost down because the load on the
condenser and reboiler will be lower. And it keeps the
capital cost down because the need for the trays inside the
column will be less, in addition to the diameter of the
column and the size of the condenser and reboiler will also
be smaller.
The total flowrate of the key components in each
individual column is constant irrespective of the sequence of
columns. While the flowrate of non-key components differs
from one sequence to the other, and that may count as a
substantial reason for the differences between the
sequences (3 > . Generally non-key components cause the
following:
1- Extra heat loads and vapour rates.
2- The gap between the temperatures in the top and the bottom
of a column will be bigger. This is so because the light
85
non-key component decreases the temperature in the
condenser, and the heavy non-key component increases the
temperature in the reboiler.
The total flowrate of non-key components can be examined
for each sequence. For the sequence in Figure 4-la:
2 m = moEB + 2 IHHPEB = 145 + 2 (7.5) = 160 kmol
For the sequence in figure 4-lb:
S m = nu>EB + mHPEB + DIEB = 145 + 7.5 + 722 = 874.5 kmol
For the sequence in figure 4-lc:
S m = ma + mHPEB = 851 + 7.5 = 858.5 kmol
For the sequence in figure 4-ld:
S m = ma + IUEB + mpEB = 851 + 722 + 152 = 1725 kmol
For the sequence in figure 4-le:
S m = mEB + 2 ms = 722 + 2 (851) = 2424 kmol
The calculations above show that the sequence chosen
(Figure 4-la) by applying the heuristics, does indeed have
the minimum flowrate of non-key components. To confirm that
86
the sequence in Figure 4-la is the best unintegrated
sequence, the energy demand for each sequence has to be
examined, and this is done in the proceeding sections.
To calculate the optimum consumption of energy in a
distillation column the optimum reflux ratio has to be found.
The lower limit of reflux ratio is set by the minimum reflux
of the column. The optimum value of reflux ratio can be found
from the relation between the reflux ratio and the number of
plates.
4.2.4 Minimun Reflux Ratio Calculations
The minimum reflux ratio is calculated by applying
Underwoods equations:
cu Xf A O.B Xf B —————— + —————— +....,...=l-q ...........(4-2)
- 8 an - e
O.A XdA O.B XdB—————— + —————— +........SR.+1 ..........(4-3)ax - 8 an - 8
To solve Underwoods equations , relative volatility data
are needed. Vapour pressure data is used instead of
volatilities, since the volatility is numerically equal to
the vapour pressure of the pure component. The Antoine
87
equation is used to calculate the vapour pressure value for
each component:
ANT(B) log P* = ANT(A) - ———————— .....................(4-4)
T - ANT(C)
The results from Antoine equation are double checked with
the values produced by using the Physical Properties Data
Service (PPDS) program, and they were found to be similar.
These results are shown in appendix (B), along with the
resultant relative volatilities which were calculated with
respect to the heaviest component in each column. An average
relative volatility was taken for each component in each
column, because the relative volatilities vary between the
top, feed, and bottom streams. These values are shown in
appendix (B), and were calculated by applying the following
equation;
aav = (CIT O.F o.B) 1/3 ............................. (4-5)
The values of q can be estimated according to the state of
the feed, as this parameter is defined by
Heat to vaporize 1 mole of the feed q = ——————————————————————————————— ..........(4-6)
Molar latent heat of the feed
After establishing all the data needed above, the first
part of Underwoods equation is solved to find the value of 0.
88
The value of 0 could be calculated by a trail and error
procedure, but for more accuracy a computer program has been
developed for this situation, the programme is given in
appendix (C). After finding the values of 0, the second part
of Underwoods equation is solved to find the value of minimum
reflux ratio. The values of minimum reflux ratio and 6 are
shown in Table 4-3. The columns that act as strippers have no
0 values.
4.2.5 Minimum Number of Plates Calculations
The minimum number of plates is calculated by applying
Fenske-s equation. The equation is given by
log (XiK / XH K>d (XHK / XLK )wSm = ———————————————————————————————————————————— ..............(4-6)
log XLK
The results of calculations are shown in appendix (D).
4.2.6 Optimum Reflux Ratio
The Erbar and Maddox correlation as shown in Figure 4-2 is
used to estimate the actual number of plates in a column
against an assumed reflux ratio value.
Therefore, the columns will be taken individually in order
89
'•CO
09C
030
070
060
0-50
0-30
0-20
0-10
0-20- - ^ -" |
I iBased on Underwood R v
——— Extrapolated
0-10-
010 020 030 040 050 060 070 080 090 100
—————— Nm /N —————
Figure 4.2 Erbar Maddox correlation.
90
to calculate the optimum reflux ratio that is compatible with
the number of plates required. This is done by multiplying
the minimum reflux ratio by different assumed values to
generate a set of values for the reflux ratio. Each value of
the reflux ratio is divided by the factor (R+l). The
calculated values along with the related curve of the Erbar
and Moddox correlation are used to generate different values
for (Sa/S). The minimum number of plates (S« ) is calculated
by using Fenske-s equation, therefore the number of plates
(S) can be calculated. These calculations are repeated for
each column, and the results are shown in tables in appendix
(D).
For each table the values of (S) are plotted against the
values of (R). The resultant curves are shown in Figure 4-3
to 4-9. From these curves the values of optimum reflux ratio
(R) can be estimated, and these values shown in Table 4-3.
4.2.7 Energy Consumption
The optimum reflux ratio values shown in Table 4-3 are
used to carry out energy balance calculations. The energy
consumption for each sequence, resulted from the energy
balance is shown in Table 4-4.
Table 4-4 indicates clearly that the sequence in
91
CO re
/-*,
z \^- w 0) -*-> o a Cu o Jumber
^.
X10
1 1
7.0
0
16.1
5^
15
.30
:
14.4
5;
13.6
0;
12
.75
^
1 1
.90:
1 1 .
05;
10.2
0:
9.3
5:
8.5
0:
7.6
5.
6 .8
0.
5.9
5.
5.1
0.
4.2
5
3.4
0
2.5
5
1 .7
0
0.8
5
0.0
0 0
1 ^—
——
——
——
——
——
——
——
-H- —
——
——
——
——
——
——
——
——
——
——
——
——
— ,1
.00
0.7
0
1.40
2.
10
2.80
3.
50
4.20
4.
90
5.6
0
6.30
7.
00
Re
Flu
x ra
tio (R
)Figure 4.3
The
relation between
reflux ratio
and
number of
plates in the
first
column of sequence (a)
in
Figure 4-1.
CO
CO
Ul
CS a (D n
xio
17.5
0
7.1
2:
6.7
5;
6.37
J6.
00:
5.63
;
5.25
:4.
88;
4.5
0:
1.1
2.
3.7
5.
3.3
7.
3.0
0.
2.6
3.
2.2
5.
I .8
8.
1 .5
0.
1.1
2.
0.7
5.
0.3
7.
0.00
0.0
0
0.9
0
1.8
0
2.7
0
3.6
0I
I |
1 1
I i
| I
I I
1 yi
• I T
T-|
I I
I
I |
I I
I I
| |
l I
r |
I I
I 1
| '
^.^w
4.50
5.
40
6.30
7.20
ReFLux ratio (R
)S
'.IO
T.O
O
Figure 4.4
The
relation be
twee
n re
flux
ra
tio
and
numb
er of
plates in
the
second column of sequence (a
) in
Figu
re 4-
1.
CO
z 1/1 0) d "o.
o_ o Numbe
100
95:
90^
85
:
80:
75:
70:
65
-
60
:
55
:
50.
45
:
40:
35:
30 ;
25:
20 15
10 5 0
• I v_0.0
0I.
50
3.00
4.50
6.
00
7.50
9.
00
Re
Flu
x ra
tio
(R)
10.5
0 12
.00
13.5
0 15
.00
Fig
ure
4.5
T
he
rela
tion
b
etw
een
refl
ux
rati
o
and
n
um
ber
of
pla
tes
in
the
thir
d
colu
mn
o
f se
qu
ence
(b
) in
F
igu
re
4-1
.
01
Z I/I Qj O a o Jumber ^_
xio
11
7.0
0
16
.15;
15
.30
:
14.4
5J
13.6
0:
I2.7
5J
1 1 .
90;
1 1
.05:
10.2
0.
9.3
5.
8.5
0.
7.6
5.
6.8
0.
5.9
5.
5.1
0.
4 .2
5
3.4
0.
2.5
5
1 .7
0
0.8
5
0.0
0
•
V^_
__
_0.00
0.70
I.40
2.10
2.80
3.50
4.20
ReFLux ratio <R
)4.90
5.60
6.30
7.00
Figure 4.6
The
rela
tion
be
twee
n re
flux
ratio
and
number of
plates in the
first
column of sequence (c
) in
Figure 4-1.
CD
05
/-\ z \_^ U1
Cb -P
o Q
.
U_
O Jumber ^.
X10
1 8
.50
8.0
7J
7.6
5;
7.23
;6
.80
;
6.3
7;
5.9
5:
5.5
3;
5.1
0;
4.6
8;
4.2
5:
3.33
;3.4
0:
2.9
8.
2.5
5:
2.1
2.
1 .7
0.
] .2
8.
0.8
5.
0.4
3
0.0
0
• \ V^_
_0.00
I .00
2.00
3.00
4.00
5.00
6.00
Reflux ratio (R)
7.00
8.00
9.00
10.00
Figure 4.
7 The relation between reflux ratio and
number of
plates in
th
e second column of sequence (c)
in
Figure 4-
1.
CO -J
2 (L 0 "a CL
O Jumber
^.
X10
1 4.2
0
3.9
9J
3.7
8:
3.57J
3.3
6:
3.1
5:
2.9
4,
2.7
3.
2.5
2.
2.31
.
2.1
0.
1 .3
3.
1 .6
8.
1 .4
7.
1 .2
6.
t .0
5.
0.8
4
0.6
3
0.4
2
0.2
1
0.0
0
I
•• — —
— i—
__
__
-,
__
__
_ ( _
__
__
__
__
__
__
__
__
__
__
__
__
__
_ i
0.00
0.
30
0.60
0.90
I .20
I .50
I .
80
2.10
ReFLux ratio (R)
2.40
2.70
3.00
X10
1Figure 4.8
The
relation between
reflux ratio
and
number of
plates in the
second column of sequence (d)
in
Figure 4-1.
oo
2 ^ £ o a u_ 0 _O £ 3 Z
X10
1 M
.2
0
10.6
4:
10.0
8:
3.5
21
8.9
6^
8.4
0:
7.8
4:
7.2
8^
6.7
2:
6.1
6:
5.6
0.
5.0
4:
4 .4
8:
3.9
2:
3 . 3
6 :
2.80
:2
.24
1 .6
8
1.1
2
0 .5
6
0.0
0
. \ ^——
-^—
—
,
0.00
1.10
2.20
3.30
4.40
5.50
6.60
7.70
8.80
9.90
ReFLux ratio
(R)
Figure 4.9
The
rela
tion
be
twee
n re
flux
ratio
and
numb
er of
plat
es in the
seco
nd column of sequence (e)
in
Figu
re 4-1.
1 .00
Sequence (a)
Sequence (b)
Sequence (c)
Sequence (d)
Sequence (e)
Column Number
Column 1 Column 2
Column 1 Column 3
Column 1 Column 2
Column 2 Column 3
Column 2 Column 3
Value of 0
21.5 2.72
21.5 3.3
2.78 1.6
31 2.72
3.5 1.6
R.
0.28 0.44
0.28 0.66
0.13 0.5
1.44 0.44
0.35 0.5
Optimum Reflux Ratio
0.36 0.57
0.36 0.86
0.25 0.6
1.7 0.57
0.45 0.6
Table 4-3 The values of 6, minimum reflux ratio, and the optimum reflux ratio for the columns in the sequences shown in Figure 4-1.
99
Heat Required in the Beboilers (MW)
Heat Rejected in the Condensers (MW)
Sequence (a)
29.855 21.062
Sequence (b)
30.472 23.125
Sequence (c)
38.433 29.27
Sequence (d)
49.352 42
Sequence (e)
52.068 49.9
Table 4-4 The energy consumption for the sequences shown in Figure 4-1.
100
Figure 4-la consumes less energy than any other sequence.
These results confirm that the sequence pointed out by
applying the heuristics, is the best unintegrated sequence,
and will be energy integrated in the context of overall
process integration.
4.2.8 Concluding Remarks
The identification of best unintegrated sequence has
reduced the energy consumption considerably, even before
taking integration into account. This appears clearly when a
comparison is taken between the chosen sequence in Figure 4-
la and the original configuration in Figure 3-2 in the
previous chapter in terms of energy consumption. The new
sequence consumes 29.9 MW as hot utilities and 21.1 MW as
cold utilities. Whereas the old configuration consumes 66.1
MW as hot utilities and 58.9 MW as cold utilities.
If capital cost is taken into account, the new sequence
has some advantages over the old configuration. The new
sequence has three columns, whereas the original
configuration has four columns. The new sequence consumes
less energy, therefore it is obvious that the condensers and
reboilers will be smaller in size. The optimum reflux ratio
has been selected in relation to the number of plates, which
means that the individual column capital cost has been kept
101
down as well.
4.3 Dehydrogenation Process
The approach adapted for the alkylation process is used
once again to select the best sequence for separation in the
dehydrogenation process. This process has four components to
be separated, these are benzene, toluene, ethylbenzene, and
styrene. The styrene will be sent to another column to be
further purified. This column is called the styrene finishing
column and is not used as a part of the sequencing problem,
this is a process requirement to insure that all heavy
components (e.g. polymerized material) are removed from the
product in the final unit. Therefore five possible sequences
can be drawn to separate the four components above, and these
are shown in Figure 4-10.
The key to the choice of best unintegrated sequence will
be dominated by the separation of ethylbenzene and styrene
(heuristic No.l). Since these two components form the most
difficult separation because of the small difference in their
boiling points which is only 9 °C. Therefore, three sequences
will be excluded from the test and these are b, d and e. This
would leave the process with two possible sequences, and
these are a and c. Both sequences separate ethylbenzene from
styrene in the absence of non-key components. By use of both
102
(a) BZ, Tol EB, Styrene
Tol
EB
EB <Styrene ^Styrene
(b)
(c)
BZJol EB, Styrene
BZ.Tol EB, Styrene
BZ
Tol EB Styrene
BZ Tol
EB Styrene
Tol
EB
Styrene
BZ
Tol
EB
Styrene
d)
(e)
BZ, Toi EB, Styrene
BZ, Tol EB, Styrene
Styrene
BZ
Tol EB
BZ Tol
EBStyrene
Tol
EB
BZ
Tol
Figure 4.10 The possible distillation sequences for separating the effluent of the dehydrogenator-
103
non-key component flowrates and energy balance calculations
over the two sequences, the sequence in Figure 4-10c is found
to be the best unintegrated sequence. The results of energy
consumed and rejected by the two sequences along with the
non-key component flowrates are shown in Table 4-5.
Therefore, the sequence in Figure 4-10c is selected to be
energy integrated in the context of the overall process
integration.
The selected sequence in Figure 4-10c found to be the same
sequence as in Figure 3-4 in the previous chapter. The
optimum reflux ratio values have been calculated following
the same procedure followed earlier in this chapter. The
related tables and graphs are shown in appendix (D).
4.3.1 Concluding Remarks
Out of five possible sequences for separating the effluent
of the dehydrogenator only two are practically valid, due to
the difficulties of separating ethylbenzene from styrene in
the presence of non-key components.
The selected sequence is found to be the same as the one
in the previous chapter, therefore the energy consumed and
released is similar.
104
Sequence (a) Sequence (c)
Energy Required By the Reboilers
(MW)
121.89 90.46
Energy Released By the Condensers
(MW)
123.2 91.53
Non-key Components Flowrate (kmol)
2406.5 679.2
Table 4-5 The energy requirements and the non-key components flowrate of sequences a and c in Figure 4-10.
105
CHAPTER FIVE
Energy Integration in the Styrene Plant
The work in this chapter is concentrated on the
maximization of energy recovery in both processes involved in
manufacturing styrene, and thus minimizing the utility
consumption in the styrene plant.
This work involves analysing the original process heat
exchanger network, setting targets for minimum hot and cold
utility consumption, designing a heat exchanger network that
is compatible with these minimum utility consumption targets,
and finally, the resultant network will be examined to seek
any improvement that can evolve the heat exchanger network to
give better energy recovery.
The two processes (alkylation and dehydrogenation) are
tackled separately, in order to accommodate the energy
integration techniques in their appropriate places.
5.1 Alkylation Process
5.1.1 Streams Extraction
The alkylation process flowsheet is shown in Figure 5-1.
Process stream data from the original material and energy
balances can be extracted and represented in a better and
106
BZ
Et_^
——
»•
Rec
ycle
d P
EB
40
'—
——
——
——
——
——
^——
——
EB
W.S
.
; ——
n
B2L
81°
^ c. E o U U8°
x^
* /
(5) r 80°
(2)
i
EB
138° fM Column
195° *-
^ y(6) y 137°
(3)
(4)
185° « c E 3 "o
o 220°
EB DEB
(7)
HPE
B ^
DEB
HPEB
(9)
EB
Resi
Fig
ure
5.1
Alk
ylat
ion
proc
ess
flow
shee
t be
fore
en
ergy
in
teg
ratio
n.
more convenient form called the grid representation, as shown
in Figure 5-2. This representation shows the actual process
heat exchanger network that is contained in the flowsheet.
The important features about this representation are:
1- The top streams are the hot streams (requiring cooling),
and these streams run from the left to the right,
2- The bottom streams are the cold streams (requiring
heating), and these streams run from the right to the
left.
3- The open circles on the streams represent the heaters (H)
or coolers (C).
4- The streams being matched (heat exchanged) can be
represented by a vertical line joining two open circles.
5- Supply and target temperatures, heat capacity flowrate,
and heat load are written as in Figure 5-2.
The alkylation process contains 9 streams (as numbered on
the process flowsheet) to be heated or cooled. The grid
representation shown in Figure 5-2 indicates clearly that
there is no heat exchange in the process whatsoever, and all
the heating and cooling duties, 29.85 and 34.403 MW
respectively, are imported from external utilities. To asses
how far away the present network is from the "best" network,
targets for minimum hot and cold utility consumption must be
established followed by designing the best network which
meets these targets.
108
Stream No.
(1)(2)(3)(4)(5).(6)
( MW, LoadlMW) Cp tT
95°+j »-/
on0 ou —————137°10C°lOD
i~>r°
X£v
^•698——— (e) ———— ^r i.36M&.065
3.19
^9.89 —————— /r\ ———
——— - 40° 4.698^ / f\° 1 OC————— *»U l-JO
——— 40° 4.065•. / n° 1 1 Q
.,_,.. , ^. on° Q onou y.oy——— te.i*y?o 11 o
0.0854 0.034 0.042
0.0229.89*i ~>
(7)(8)(9)
196*————— 221°—————
OH = 29.85MW
'17.1
'11.3
1.45
U8 195°
= 34-4MW
17.111.3
17.111.3
Figure 5.2 The grid representation of the heat exchanger network contained in the alkylation process flowsheet before energy integration.
109
Targets can be set by the set of calculations that form
the problem table or by constructing the composite curves.
The problem table is easier and quite adequate for giving
both the targets and the pinch location. The composite
curves illustrate well the heat flows within the process.
Therefore, both procedures will be used to consolidate the
integration results.
5.1.2 Targetting and Related Design
The problem table for the alkylation process has been
derived for a minimum approach temperature (^Tmin) of 10 °C to
generate the heat flow cascade shown in Table 5-1 . The value
of Tm i n is usually chosen by engineers from their previous
experience to offer a good tradeoff between energy
consumption and capital cost. Table 5-1 reveals that the
minimum hot utility is 29.3 MW (1.98% less than the original
usage), the minimum cold utility is 33.8 MW (1.75% less than
the original usage), and the pinch is located at a
temperature interval of 153 "C. A Pinch location at 153 °C
means that the hot streams are pinched at 158 °C and the cold
streams are pinched at 148 °C, since the minimum approach
temperature was taken to be 10 °C.
The importance of the pinch can be seen in designing the
heat exchanger network that maintains the targets given
110
Interval Temperature (°C)
226
225
201
200
180
154
153
133
132
90
76
75
35
Heat Flow (MW)
29.256
27.806
27.806
16.506
16.506
17.078
0.000
0.44
11.662
14.35
16.442
26.481
33.817
Table 5-1 The Heat flow cascade of alkylation process before the process being improved.
Ill
Stream No.
(1 )
(5)
(6)
(4) 185° ————————
(31
(2)
17.1/Q 11Qfi°— ————— ffb ————1 O i ' «''-' x_J/
11.3 ( 9 ) 221° ———— ® ———
1.45
(10)
Qxp =0.594
QH = 29.85 QHmfn =29.25
158° { MW j AH(MW) Cp *C
J95 ———— © ———— 40° 4.698 0.0854 4.698
1 81° fT\ -• '-»• ftn° Q AQ Q fto
I 9.89 J1382 ———— © ——— 137 ° 11.2 11.2I 11.2
® _ / O *^ 4 O f\ f\ O^l—————————— »" 40 3.19 0-022 3.2
I n i ° /PN ^ / n° / HKC: nn/ojM/ ————— ^ ——— *• t»U 4.UDD U.U^^
1 4.065i
80° ———— © ———— 40° 1.36 0.0341.36
I I
U8 I7.1 17. 1
i/~»r~O1 95 n.3 n .3I
*"io r\°220 1.45 1.45
148° Qc = 34.4
Figure 5.3 Afky(ation process heat exchanger network with coolfng duty above the pinch (energy targetsare not maintained).
113
Stream No.
(5)
(6)
)185
(2)
(3)
(7)149'
(9)221'
11.3
1.45
e-
158'
l81c
16506 0.594e—'us*
22 Oc
4.698
9.89
11.2
2.596
4.065
1.36
.MWAK(MW) Cp ( *C ]
0.0854
40C
- 80° 9.89 9.89
137 11.2 11.2
40° 3.19 0.022
40° 4.065 0.042
0.034
17.1 17.1
11.3 11.3
1.45 1.45
148 C
= 0Cmin = 33.8
Figure 5.4 Alkylation process heat exchanger network that reaches the process energy targets.
114
36
-
= 2
8
Cold
U
tility
"S32
'5 cr 91 H
ot
Util
ity
200
18
36
54
72
90M
inim
um a
ppro
ach
tem
pera
ture
(°C
)
Fig
ure
5. 5
A
lkyl
atio
n
proc
ess
utilit
y
requ
irem
ents
fo
r a
rang
e of
min
imum
appr
oach
te
mpe
ratu
res.
streams run above the pinch. In this case the process needs
to be improved further, so as to draw more streams above or
below the pinch in order to increase the heat recovery via
inter process exchange.
5.1.3 Energy Saving Techniques and Process Improvement
The composite curves corresponding to the process are
shown in Figure 5-6. This figure indicates that the
composite curves are dominated by the reboilers and
condensers of the distillation columns. These distillation
columns are not appropriately placed, as they span the pinch.
Before any attempt is made to integrate the distillation
columns within the process, energy recovery by inspection
will be sought within the process to find out whether or not
the process would benefit from such a scheme.
5.1.3.1 Energy Recovery By Inspection
The feed stream to the distillation process appears to be
the only stream that can benefit from the heat available in
the process. This stream was not included before in the
original network design, and is available at a temperature of
40 °C. So the scheme now is to raise the feed temperature
116
300
246
-
E 01
Q. E
138 30
^Min
imum
ho
t u
tility
= 2
9.25
MW
ATm
in=1
0CQ
Min
imum
co
ld u
tility
= 3
3.81
MW
L _
__
__
__
I__
__
__
__
_L
r
I__
__
__
__
_I_
__
_
i
0u
28E
nth
alp
y (M
W)
5670
Fig
ure
5.
6 A
lkyl
atio
n
proc
ess
com
posi
te
curv
es
befo
re
any
impr
ovem
ents
.
from 40 °C to its bubble point, which is 99 °C. By doing this
a new cold stream is created. This feed stream is now
designated stream 10 in Figure 5-7.
Raising the feed temperature would increase the reflux
ratio at the first column due to the change in average
relative volatilities. Therefore new values are used (these
are calculated as in the previous chapter) to calculate the
new energy balance around the first column.
This technique would reduce the hot and cold utilities
considerably as seen by the process heat flow cascade in
Table 5-2. Table 5-2 shows that the hot utility target is
25.2 MW (15.7% less than the original usage), and the cold
utility target is 29.7 MW (13.7% less than the original
usage). The pinch location is at interval temperature of 153
°C, which is the same position as it was before ( i.e has not
changed), because the feed stream (stream No,10) was
introduced with both its supply and target temperatures below
the pinch. Thus, the energy targets are varied but not the
pinch location.
The reduction in hot and cold utilities occured because
heating up the feed stream replaces some of the cold
utilities, since it has been achieved through process to
process heat exchange. By heating up the feed stream, the
heat load on the first distillation column reboiler will be
118
158 MWStream No. AH(MW) Cp l C
(1) J95 ————— © ——— -40° 4.6981 4.698
(5) |81- ———— © ——— 80° 10.9 | 10.9
IS~\ 1(4) ————————— 0 ——— 1 ———
/ o \ nn° - -tI £. I OU
\ (31 il37°\ «J J • I ^ /
I
( 7 ] 1492—— ® ————— 0 ——— 148° 12.406 0.594 |
i o ] 196°" ® ———————— 19^ 131 11.3 1
I ( 9 ) 221°" ® ———————— 220,°1.45 ,
, , , lorVL r
J \^ ~~ I0/ 1 I.Z
6.07
—— © ——— 40° 3.192.596
© .. M. / r>° 1 TC^ t»u I.JD 1.36
—— © ——— - 40° 4.0654.065
13
11-3
1.45
s\ /n° e; n
0.0854
10-9
11.2
0.022
0.034
0.042
T3
11.3
1.45
n n«5.13
148
= QHmm=25.156MW = 29.69MW
Figure 5.7 The network design of alkylation processwhen the -feed to the separation process is heatedup to its bubble point by using one unit.
119
Interval Temperature ( ° C )
226
225
201
200
180
154
153
133
132
104
90
76
75
45
35
Heat Flow (MW)
25.156
23.706
23.706
12.406
12.406
12.978
0.000
0.44
11.662
13.454
13.132
14
24.968
27.86
29.69
Table 5-2 The heat flow cascade of alkylation process after heating up the feed to separation process to its bubble point.
120
reduced, thus the hot utilities will also be reduced.
Different network designs can be generated for this
situation that maintain the targets above. As the region
above the pinch has not been effected, then the network
design for this region would not change, therefore the
changes are restricted to the region below the pinch. Some of
these designs are selected for analysis, and are shown in
Figures 5-7, 8, 9, 10, and 11.
5.1.3.1.1 Analysis of the Network Design
The design should be carried out away from the pinch, as
the pinch indicates the most constrained point in terms of
driving force. Below the pinch, a hot stream heat capacity
flowrate (Cp) has to be greater than that of the cold stream
it is to be matched with, but this constraint can only be
considered at the pinch, away from the pinch there is no heat
capacity flowrate (Cp ) constraint* 75 > . Therefore, different
designs may be derived to satisfy the heat demand for stream
No. 10 (feed stream), as it is the only cold stream existing
below the pinch, thus satisfying the targets (these designs
are shown in Figures 5-7, 8, 9, 10, and 11).
The network design shown in Figure 5-7 matches stream
No. 10 with stream No. 6, and the heat load is maximized so
121
Stream No. 158 AH(MW) Cp Ci OI.O — ~ -
1
ir\ Ifil0 {T\\ J ) 01 V^y
1 10.9
] 11.2 (41 18*"° C~*) (&
1 2.596
(2) .80° ———— ©- 1.36
(3) 1372 ———— P^\ -J 1 | U / V
1 1 1
1 -7 \ \/tf« (Eh I) Jl/C°l // it»y* \c/ v^ I4o 12.406 0-594 [
(8) 196° ——— ® ————— 195°
11.3 1(9) 221°- —— @ ————— 220°!
1.45
^
$
pr<57-i,U 4.Dyo U.U0^40.855
—— -80° 10.9 10.9
—— M37° 11.2 11-2
—— *" 40° 3 19 0 022•»W W . 1 9 W.Wfcfc
—— 40° 1.36 0.034
-©«-400 4.065 0.0421.797
13 13
11-3 11-3
1.45 1.45
(10) 199- O w **u o-u u.uo/ 1 2-268 2-862 1 1 1 148°
=25.156MW = QCmin = 29.69 MW
Figure 5-8 Alternative network design for alkylation process when the feed to the separation process is heated up to its bubble point by using two units.
122
Stream No.
(1)
(5)
(6)
(0185° ———————— f
(2)
(3)
(7)149°-*— © ————— G ———— 'i<to
158° I AH(MW) Cp ( t }
95 (Q^ • **u 4».uyo u.uo:>*« | 4.698'81° ———— © ———— 80° 10.9 10.9| ^0.91 0 'l*3O /P\ ~ « "> »•» O 44 ^ 4 4 "»,1oo ———— fe^ ——— *-137 11.2 11.2
11 ? 1 93°
"\ | (^ ^ — ' n° i I" n ror>
1 0
1
1
'137°137
11
^ L , «o12.406 0.594 '
fni1Qp°— r (Hi —— 1QTloJiyo vn/ iyj i11.3
(9)221° — @ ——————— °'>n0 1.45
(10)
1 1
1.166
—— © ——— - 40° 1.36 0.034 1-36
/go lATminrS"
ji/ V^ •" t»U «*.UUJ U.U*4^
0365
13 13
11.3 11.3
1.45 1.45
'99 1— O VN^ **U J.IO U-UO/
1 1.43 3.7 I
148
Q H = Q Hmm=25.156MW Q C = = 29-69MW
Figure 5-9 Alternative network design for alkylation process when the feed to the separation process is heated toits bubble point, by using two units andtemperature violation.
123
, MW I «xStream No.
(1)
(5)
(6)
(2)
(3)
(7) 149°- ——— @ ——— b —— •«•«
iV° Ci SIloi° fr\Ol •- • —— tjs? —1 10.9
] 11.2
1 0
1 1137°
I I I
K d/ 00
12.406 0594 I(8) 196° ———— @ ————— 195°,
11.3 ]O O • »-i» -.-.-i rt*\ O^^\(9) 2^1— ———— @ —————— oon '
1.45
(10)
I
J v^ 1.166
1.36 536?
^
42.1° S ?
o AH(MW)AD 'w^V / r\O i e noKfcjr 40 4.o9o 4.498
—^80° 10.9
»n7° 11 ?^^ I O / ll«^
..^r /n° T iq^^ *»u o- 1 y
^_ / n i *3 R—— ""U.U 1 • JO
-©-40° 4.0650.565
13
11.3
1.45
QQO O W ^-^ *»U J-U
j 1.43 3-5-0-2
I I 148°
Cp"C '
0.0854
10.9
11.2
a022
0.034
0.042
13
11.3
1.45
0.087
°H =QHmin = 25.156MW = 29.69MW
Ffgure 5.11 Alternative network design for alkylation processwhen the feed to the separation process is heated up to its bubble point, by using three units-
125
that stream No.10 is ticked off (eliminated). The rest of the
heat load on stream No.6 is satisfied by cold utility, as are
the other hot streams below the pinch. In this design, only
one heat exchanger is needed to satisfy stream No.10. If
stream No. 6 should not be matched with stream No.10 due to
the need for the heat load on stream No. 6 elsewhere in the
process or for any other reasons, then alternative networks
should be sought.
Stream No. 10 can not be ticked off by heat exchange with
any other stream using just one unit (heat exchanger), since
the heat load on any other individual hot stream can not
satisfy the heat demand of stream No.10. Therefore, if
another design is sought, then more than one stream is
required to be matched with stream No.10, thus using more
than one unit.
The designs in Figures 5-8, and 9 show that stream No. 10
is satisfied by using two units (heat exchangers). In Figure
5-8 stream No.10 is matched with stream No.3 first, and the
heat exchange is maximized to the limit, until the streams
are at the minimum temperature difference of 10 °C. This
match results in a temperature of 73 °C in stream No. 10 and
83 "C on stream No. 3. The rest of the heat load on stream
No.3 will be dealt with by cold utility. The second match is
with stream No. 1, this match will satisfy the rest of the
heat demand for stream No.10, and reduce the temperature on
126
stream No.l to 61.5 °C. The rest of the heat load on stream
No.l is dealt with by cold utility, the same as for the rest
of the hot streams.
In the case shown in Figure 5-9 stream No.10 exchanges
heat first with stream No.4, using all of this streams
available heat would result in a temperature of 93 °C on
stream No. 4 (which is the lower limit for stream No.3) and
82.4"C on stream No.10. The second match on stream No.10 is
with stream No.3. If the heat load on this match is maximized
to satisfy the rest of heat demand on stream No. 10, then the
minimum approach temperature will be violated by 1 °C.
Specifically the temperatures are 49 °C on stream No.3 and 40
°C on stream No.10. To manipulate this violation, more heat
exchangers need to be used, but instead, stream splitting can
be introduced to keep the number of units the same, and at
the same time offering more flexibility by generating bigger
driving forces.
Figure 5-10 displays the case where stream No.10 is split.
The two branches resulting from the split have different heat
capacity flowrates. These are chosen to satisfy each branch
demand, and the values of heat capacity flowrate must not be
less than the minimum and must not exceed the heat capacity
flowrate of the main stream, otherwise the network is
infeasible (76 >. The minimum heat capacity flowrate for branch
a (Cpa)min may be calculated as follows;
127
Qa(Cpa )min = ——~~——~————-
Ta
where Ta - Maximum temperate that branch a can reachwithout violating the minimum approach temperature - The initial temperature of branch a
2(Cpa )min = ————————————————————— = 0.0185
( 158 - 10 ) - 40
The minimum heat capacity flowrate for branch b (Cpb)«in
may be calculated as follows;
3.13(Cpb)min = ———————————————————— = 0.0355
( 138 - 10 ) - 40
(Cpa)nin + (Cpb)min = 0.0185 + 0.0355 = 0.054
The result of the minimum heat capacity flowrates of the
two branches is less than that of the main stream (0.087),
therefore the network is feasible. Now branch a is matched
with stream No.4, and branch b matched with stream No.3. The
temperatures resulting from these matches are 67.1°C on
stream No.4, and 62.5°C on stream No.3. Thus there is no
minimum approach temperature violation, and there is no need
for using more heat exchangers.
The following design will deal with the problem of the
minimum approach temperature violation by increasing the
number of heat exchangers as shown in Figure 5-11. In this
network design, three heat exchangers have been used to
satisfy the heat demand of stream No.10. This technique is
128
not advisable since it uses more units than the other
designs, and the more units the design uses the more
difficult the process control. Also the capital cost is
likely to increase being strongly influenced by the number of
units used.
Of the designs shown (Figures 5-7, 8, 9, 10, and 11), the
design in Figure 5-7 may be considered as the best, as it
uses the minimum number of units, and it offers the most
independent control. Notice that the number of units in the
other designs was more than the minimum because there was no
scope for the tick off rule to be applied. If stream No. 6 in
Figure 5-7 can not be matched with stream No.10, as stated
earlier, then the design in Figure 5-10 may be favoured,
since it offers greater flexibility in terms of driving
force, and this would lead to a smaller heat exchange area as
a consequence of the increase in driving force.
However, the best network design cannot be chosen until
all possible process improvements have been examined.
Therefore the process will be searched again to find out
whether or not there are some further improvements that may
be made. This time process parameters will be taken into
account, since all possible heat exchange probabilities for
the existing situation have been examined above.
129
5.1.3.2 Increasing Energy Recovery by Process Improvement
The composite curves for the process including heating up
the feed stream is shown in Figure 5-12. The most important
feature of these curves is that the hot utility is needed
mainly for the distillation columns reboilers. Since the
temperature level of these reboilers prohibits their heat
loads being met by any other process stream via heat
exchange. Therefore, the only scheme that can make the best
use of available energy is to shift the cold composite curve
below the hot composite curve, or to shift the hot composite
curve above the cold composite curve. This shifting will
change the pinch location, also the position of the streams
relative to the pinch. This would create more opportunities
for heat exchange.
The objective of this shifting, is to change the position
of the distillation columns relative to the pinch, and place
as many columns as possible completely above or completely
below the pinch. So that one or more reboilers can be met by
other process streams. This technique increases the heat
recovery and decreases the utility demand.
Analyse of the composite curves in Figure 5-12 shows that,
above the pinch, the closest stream is the first column
reboiler, and below the pinch, the closest stream is the
second column condenser. Therefore, by lowering the reboiler
130
300
246
£192
CL E13
8
30
ATm
in 1
0°C
HI,
Min
imum
col
d utilit
y
2969
Min
imum
ho
t u
tility
*
25.1
5 r
014
15
42
Enth
alp
y (M
W)
5670
Fig
ure
5-
12A
lkyl
atio
n pr
oces
s co
mpo
site
cu
rves
af
ter
heat
ing
up
the
feed
to
the
se
para
tion
proc
ess
to
its
bubb
le
poin
t.
temperature, and raising the condenser temperature, the
composite curves will fit comfortabty on each other. This
would result in a complete distillation column above the
pinch (second distillation column), since the pinch location
is changed. This format would allow the heat to be
transferred from the second distillation column condenser to
satisfy the heat demand of the first distillation column
reboiler.
This changing of the condenser temperature of the second
column, and reboiler temperature of the first column is
obtained by varying the operating pressures for these
columns. In the first column the pressure is reduced to 0.9
bar at the bottom, and for the second column the pressure is
increased to 1.4 bar at the top. These pressures are chosen
to make the difference between the condenser temperature and
the reboiler temperature 10 °C, as this is the minimum
temperature difference allowed in the design. Therefore, the
condenser temperature will be 148 °C and the reboiler
temperature will be 138 °C.
The change in relative volatilities, reflux ratios, latent
heat of condensation, and heat loads due to the variation of
operating pressures is taken into account in calculating the
energy balances around the specified distillation columns.
The process heat flow cascade, after changing pressures,
132
is shown in Table 5-3, as are the energy targets and the
pinch location. The design targets now are 14.25 MW for hot
utility (52.26% less than the original usage) and 18.8 MW for
cold utility (45.35% less than the original usage). The pinch
location is at an interval temperature of 142°C. The
composite curves that portray this situation are shown in
Figure 5-13.
The corresponding network design (that maintain the
targets above) is shown in Figure 5-14. The important feature
in this network is the part above the pinch. Once the best
design above the pinch is decided, then it is easy to
complete the whole network because most designs below the
pinch have been studied in the previous section.
Only the one design option considered above can be derived
in the area above the pinch as shown in Figure 5-14. This
design matches stream No. 7 first with stream No. 6 and then
with stream No.4, and the rest of the heat demand is supplied
by hot utility, as are the other cold streams. This design
offers the minimum number of units, since the tick off rule
can be applied on the two matches. However, this design would
result in three exchangers on stream No. 7 (first column
reboiler). This would make the first column difficult to
start up, shut down, and control. Therefore this design is
not desirable. If less exchangers are used for stream No. 7,
then energy targets will be violated. Therefore, the process
133
Interval Temperature (°C)
226
225
205
204
180
143
142
104
90
68
67
45
35
Heat Flow (MW)
14.25
12.804
12.804
0.424
0.424
1.238
0.000
2.47
2.162
3.556
14.84
17
18.8
Table 5-3 The heat flow cascade of alkylation processafter pressure changes in distillation columns
134
300
CO 01
246
o o
192
a e 138
AT
min
10
°C],
Min
imum
^C
old
Util
ity-1
8.36
MW
30
Min
imum
Hot
Util
ity*
14.2
5 M
W
012
24
36
Ent
halp
y(M
W)
Fig
ure
5.
13
r
4860
Alk
yla
tion
pr
oces
s co
mpo
site
cu
rves
a
fte
r ch
angi
ng t
he p
ress
ure
to
mak
e th
e te
mpe
ratu
re
diff
ere
nce
be
twee
n fir
st c
olum
n re
bo
iler
and
seco
nd
colu
mn
cond
ense
r 10
°C.
Stream No.
ID
(5)
(6) 148- ————————— es
(2)
(3)
f 7 ] nfl°_ /uV r
j
^ S
147*
|^0 ^^
MW AH(MW) CpCC
_ /rto|95 " (jj)' — fciVJ *»-U3U U.UOJt*
, 4.698(73° /^\ _ T0° 110 11 0
I
I
I
I
i SI 72°I I
Q
I
I
S In -701 l\ loo •• \r\J \_? \_; 1 '"•"0,424 0.836 11.6
(83 200°-~® ———————— 19$. 12.38
( 9 ) 22 f~@ ———————— ™°1.45
(10)
1 J
11.2
11.16 11.16
"^ j^» / n° *$ 1^1 no*?*?J ———— \S57^ '*'-' J I-* U.U^^
0.354
— © —— -40 1.12 0.035 1.12 722°S~\ i /P\ - / nr /co nn/o
K
1 ^ ———— fc^-t
1.5
12.42 12-42
1238 1238
1.45 1-45 C Pa = 0.0338
^>— £ 0° 5.13 0.087s >rb
| 3.13 Cpb= 0.0532
OH = in = 14.25 MW137'
= OCmin = 18.8MW
Figure 5. 14 The network design for alkylation process when the pressure is changed to make the difference between the second column condenser and first column reboiler 10°C.
136
has to be examined again to find out whether or not there are
any other advantageous improvements that can give a better
network design.
5.1.3.2.1 Process Examination for More Improvement
If the number of units on stream No.7 in Figure 5-14
is reduced, then either the heater is removed or the match
with stream No. 4 is deleted. If the heater is removed then
stream No.7 would not reach the target temperature, and would
violate the energy targets. If the match with stream No.4 is
deleted then stream No. 4 will be cooled down to the pinch
temperature by using cold utilities above the pinch, and this
is forbidden.
The only successful approach to this problem is to
transfer streams No.4, 6, and 7 below the pinch and seek a
match with the heat load that can compensate for the two
units mentioned above. The three streams cannot be
transferred below the pinch unless the pinch location is at
least at an interval temperature of 185 °C, since the highest
temperature of these three streams is 185 °C, i.e. the
supply temperature of stream No.4.
The solution is to modify the composite curves in Figure
5-13. This is done by increasing the driving force between
137
the two curves, so as to shift the cold composite closer to
the hot composite and move the pinch to an interval
temperature of 180 °C. This is achieved by increasing the
temperature of the second column condenser and decreasing the
temperature on the first column reboiler to give a
temperature difference of 20 °C. The resultant temperatures
and pressures are 153 °C and 1.6 bar at the top of second
column, and 133 °C and 0.8 bar at the bottom of first column.
The composite curves for this situation are displayed in
Figure 5-15, and different network designs may be derived for
this situation. The network shown in Figure 5-16 is chosen
to be the best, since it is the only design that can satisfy
the need of stream No. 7 using only two units. The best
network design (Figure 5-16) is compared with the previous
design (Figure 5-14), and the following results obtained;
1- The design in Figure 5-16 requires 11 units while the one
in Figure 5-14 requires 12 units. Therefore, the capital
cost is reduced.
2- In Figure 5-16 the temperature difference between streams
No.6 and 7 is 20 °C, while in Figure 5-14 it is 10 °C.
This increase in the driving force saves capital cost with
no great impact on the utility consumption, as the cold
utility stays the same and the hot utility is increased
by only 0.15 MW.
138
CO
CD
246
o o <U19
2
Q Q. 13
8 30
Min
imum
H
ot U
tility
r
Min
imum
X
Co
ld U
tility
=16
.84
MW
012
24
36E
ntha
lpy
(MW
)48
60
Fig
ure
5.15
A
lkyl
atio
n
proc
ess
com
posi
te
curv
es a
fte
r ch
angi
ng
the
pres
sure
to
mak
e th
e te
mpe
ratu
re
diffe
renc
e be
twee
n th
e fir
st c
olum
n re
bo
iler
and
seco
nd
colu
mn
cond
ense
r 20
°C.
1<
Stream No.
(11
(53t %r 1
(6)
(4)
(2)
(3)
(8) 208- — 0 ——— 207°1235
(9) 2212, — Q ——— 220°U5
(10)
35°
1
95°J «•/
69°
153°
4 ft J™O
Cfl°DO
152° fi ji. r
3-!
-e
3-
-eo.<
>26
113
— c
70°
W9i 11
>8.5C
QC
-€
5-
3-.13
-4i.i
= c
MW AH(MW) Cptec '
76-2* ^^£W/.n° L ^Qft n nftQA.7 ^^ *t U «*,J3O U>UO9^*
3.094
__ cO° 1 1 o 11 "5———— *-DO 11. J 1 I.J
•.i 1^^5 11 n 11 n
©— / n° o i o n no o^**»u J.iy u-uz/ 2.25
© -40° 0.95 0.0340-95
©/ n° / fi 1 n n / *3— -uu t*-.O I U.Ut»J 1.28^
132° 1207 1207!>«/£. * b* W / * fc*\»/ /
12.95 12.95
1.45 1.45
5 ——— 40° 5.13 0.0875CH
'Cmin = 18-8
175 (
Figure 5-16 The network design for alkylati'on after changmgthe pressures to make the difference betweenfirst column reboiler and second column condenser 20°C
140
3- The design in Figure 5-16 requires no stream splitting,
and stream No.7 is satisfied by only two units. Therefore,
this design offers more flexibility and better control.
The process at this stage is well integrated, since it
offers no more scope for integration. The improved process
flowsheet is shown in Figure 5-17.
5.1.4 Process Utility Levels
The integrated process grand composite curve shown in
Figure 5-18 reveals that, above and below the pinch utility
can be introduced at one level only. Above the pinch the
utility used is steam, and can be introduced to the process
at temperature and pressure of 231 °C and 30 bars
respectively. Whereas below the pinch, the utility used is
cold water, and can be introduced at a normal temperature
around 30 °C.
5.1.5 Concluding Remarks
Process energy integration on the alkylation process after
selecting the right separation sequence has resulted in
considerable savings in both hot and cold utilities. The hot
and cold utility requirements in the alkylation process
141
Stream locationi_
JC
en .*QJ"o
£O
<*—
(/) CO 02:
C 6 H 6
C 2 H^
C8 H, 0
CIO H U
C) 2 Hi8
Total
A51540.25
————
51540.25
B
25421
25421
C
19164-8
1133-143
20297.943
D
66477.75
76629.526
19484.858
1145-863
163737.997
iJ
101046-18
101046-18
Jl
66477.75
66477.75
J2
34568.43
34568.43
K
122607.24
512
123119.24
Ki
76629.526
320
76949.526
K 2
45977.715
192
46169.715
M
76629.529
19484.858
1145.863
97260-25
N
19164.858
1145.863
20310.721
P
12-72
12.72
A,
118018
118018
R
18014-966
1087-817
19102.783
S
1149.89
67.988
1217.878
X
19164.858
1133.143
20298
Alkylator Was hi ng system Benzene column Ethylbenzene column Polyethylbenzene column
Benzene
Ethylene
Ai I1.3 bar
H a JDin c^o
J2
-»~ Ethylbenzene
K
aJQ
UD
K,
N
Stream locationHeat content with reference temperature of 25& pressure of 1bar
Temperature (°C)
A
0-377
40
A]
0.8637
40
B
0-166
40
C
0-149
40
D
5.9t
95
E
4.388
77-2
F
1.21
40
G
1.21
40
H
2.732
57.6
I
6.3
99
J
13.546
69
Jl
1.4657
68
J2
0.762
68
J 3
0-486
40
K
19-758
153
Kl
5-439
152
K 2
3-163
152
K 3
1.87
70
K 4
0.579
40
L
3.39
185
M
5.67
134
N
2-1
208
P
0-001
221
R
2.369
185
S
1.021
185
T
1.2712
185
X
2.292
185
$-11 Alkylation process flowsheet after being energy integrated
300
246
192
o o
132
Q. e
30
^Hot
Util
ity 1
4.4M
W o
t te
mpe
ratu
re o
f 23
1 °C
Cold
U
tility
18.8
MW
at
tem
pera
ture
of
3Q°C
08
12
Ent
halp
y (M
W)
1620
Fig
ure
5. 1
8 Th
e gr
and
com
posi
te
curv
e of
alk
ylatio
n p
roce
ss
afte
r be
ing
ener
gy
inte
gra
ted
.
before the energy integration carried out in this chapter are
respectively 29.85 and 34.4 MW. Due to the integration of
energy the hot and cold utility requirements have become 14.4
MW (51.8% saving) and 18.8 MW (45% saving) respectively.
5.2 Dehydrogenation Process
5.2.1 Streams Extraction
The dehydrogenation process flowsheet shown in Figure 5-19
contains 16 streams (as numbered on the flowsheet) to be
heated or cooled. The network that is involved in the process
flowsheet is represented as a grid in Figure 5-20. This grid
representation reveals that there is some process to process
heat exchange as shown by the two heat exchangers for streams
Nos.l and 15, and Nos.l and 16. The rest of the process
streams are left to be satisfied by external utilities. These
external utilities amount to 503.086 MW and 224.945 MW as
cold and hot utilities respectively.
5.2.2 Targetting and the Old Design Failures
The integration will be initiated by setting energy
targets, in order to find out how much energy saving this
process can offer. This will be followed by inspection of the
main "failures" of the old design that prevented it from
144
Stea
m16
0
Stea
m g
ases
cr
ude
styr
ene
A 4J
O O C if71
0
56^
0 po x» £ 01
. -J
EB f
rom
alk
ylat
ion
proc
ess
U)
1 R
ecyc
led
EB
Styr
ene
EB T
ar
01
Fig
ure
5.19
D
ehyd
roge
natio
n pr
oces
s flo
wsh
eet
befo
re
ener
gy
inte
grat
ion.
Streom No.
(1)(2)(5)(6)(7)(8) (U)
(15)(16) (3) (9)(13)(11)
(12)no)(4)
( MW, AH(MW) Cp °C
~o 431.6 V,2«4.3^+AJ+J V
105°*\j*J
57°+J 1
fftOJO
^*_o57 ———————
107° ——————
>» V.
k520° —————— & ——— •71 n°. ^ 52>56 r
[^ ^70.6
321.02
4.63VS/86.403
/^15.876
^0.506
1.23
s
•105°
•104°
__o• 57
• 79°7L°
4 f" f\
83.9- ai-o 385
/i. "• ——
110° ——** O75 - —
117° —— 107° ——
r>6.89
1^1.336
^15-225
^0.583
^82-987
————————— i»3
—————— 96°
————— 97°
11F°MDme®——————— IUO
C7«>28.30^
181.24321.02
4.6386X0315.8760-5021.23
52.56141.92.4346.891.33615.2250.58382-98728-304
0.39321.024-63
86-40315.8760.5020,037
0.146
0.2580.0976.890.10315.2250-58382.980-275
Q H = 221.6MW = 500.2MW
Figure 5. 20 The grid representation for the heat exchanger networkcontained in the original dehydrogenation process flowsheet.
146
reaching these targets. After that the heat exchanger network
design that is compatible with the energy targets will be
sought.
The results of the problem table (taken for the assumed
minimum approach temperature of 10 °C) as shown in Table 5-4
indicate that, the process is pinched at interval temperature
of 101 °C, and the minimum hot and cold utility targets that
have to be consumed by the process are 123 MW (44.5% less
than the original usage), and 401.6 MW (19.7% less than the
original usage) respectively. The utility targets show that
the process can be run with much less energy consumption than
the original used. The reduction in cold utility due to
setting targets is always equal to the reduction in hot
utility, because the heat recovery is done by interprocess
heat exchange. This energy saving accounts to 98.7 MW.
The main reason for this great difference between the
original process energy consumption and the energy targets
is, the usage of heating and cooling duties at the wrong
places. This has resulted in a situation where most
interprocess heat exchange is not exploited. This fact can be
shown very clearly when the process pinch location for
minimum approach temperature of 10 °C is applied to the
original network design as illustrated in Figure 5-21.
Figure 5-21 shows that heating and cooling duties have
147
Interval Temperature (°C)
715
560
525
165
122
121
115
112
111
102
101
100
99
80
79
75
74
69
62
54
53
52
51
Heat Flow (MW)
123.5
83.5
88.2
84.2
89.3
88.8
89.5
89.5
6.6
6.7
0.0
0.2
320.9
316.4
300.9
299.6
299.8
298.1
295.5
294.7
294.7
381,1
401.6
Table 5-4, The problem table for the dehydrogenation process before any improvements to the process, derived for minimum approach temperature of 10 'C.
148
106 -MWream No- I
_ .^«AO *) 0 / ?^ ^^*r-° S~\ *Ol-Q X"\ /O*tO /2w\(1) 565 ———— © —— P-^ —— ——————
(2) (5)
(6) (7)(8) lU) 107°
(15) 520° • ffij6 ———o oc°itet 7in°- <Th 385 r
X VS? 70JB. _. _ oI1052 —————
1 o ^ibo
o ae.^o:| ono 1S87
L 0.502
H-23
160J
AH(MW) Cpv t__ . ir»c° ifli r\K / n TO/— ——— *~ lUb IOI.UOH u.jy*«
© ——— -104° 321.02 321027S\ CC° / CO / CO{JQ ————— - bo 4.DJ 4>.DJ
©__ r-*?® DC / no ftc / n*?———— *• b/ ob.*»UJ oD.t»Uo
© CC° 1C D*7C 1C O"7C———— *• bo ib.o/o ib.o/o ^ ——— 79° 0-502 0.502
__ 7/° 1 90 n m7———————— »•/*» I.ZJ U«UJ/
52.56 0.146 i/ 1 a n OQft
(3) (9)(13) IHf- lll)(12) H7°-
(10) (O
83-9 58
J.89
1.336
0.583
82'987
2.434
•74
-116C 15-225
28-3
96'
2-434 0.097
6-89 6.891.336 0-10315.225 15.2250-583 0.58382-987 82-987
= 221.6MW*Q Hmln Q C = 500.2MW* Q Cmin
Figure 5.21 Dehydrogenation process heat exchanger networkwhen pinch locatfonistaken into account. This is before any development is done on the process.
149
been carried out below and above the pinch respectively.
Below the pinch, heating duties amounting to 28.4 MW have
been used, as indicated by dark circles on streams No.3, 4
and 11. These heating duties must not be used, and the heat
demand of these streams should be satisfied by matching them
with hot streams below the pinch. By doing this, 28.4 MW will
be saved from both hot and cold utilities.
Above the pinch, cooling duties amounting to 70.3 MW have
been used, as indicated by black circles on streams No.l and
14. These cooling duties must not be used, and the heat load
on these streams should be given to cold streams above the
pinch. By doing this 70.3 MW will be saved from both hot and
cold utilities. Therefore, the summation of the amounts of
energy saved below and above the pinch will be 98.7 MW. These
98.7 MW are the difference between the original usage of
energy and the minimum targeted to be consumed by the
process.
5.2.2.1 The Design That Reaches the Energy Targets
The heat exchanger network design that maintains the
energy targets above is shown in Figure 5-22. The most
important feature about this design is that stream No.16 is
heated up by process to process heat exchange to a
temperature of 555 °C instead of 385 °C (the old maintained
150
106 MW
(1
(5) 16)
(7)(8) IK)
(15)(16) (3) (9)
(13)(11)
(12) (10)
Cpa =0^6 173° a |occ° !PQ * r in?12*5 123.5° '
565° —— < V
Cpa s0.134
107° ————
_ ^»«»O f
J
*N520 ———— & —o 52.56
i7 ^\lyV^N!
172.7°
^710° —— ® g" 0 6 ————
40 101.9
_ _o97 — —————————
1 1 U vj'1.336
1 1 / Vtx0 a 58 3107 ^ YD) —— — — — -i ux vty
81 r1
^CJr^ 1 ^X^ V
-( 64-9
-11
w-€^1-97 } ———
1 V
** M A|57°^1 r opI 58| r» ^»O57 ~
IOU
-e— ' ——
>397°
0
6
^
j •
1 160°160°
1 7/0i 74-*- ^96°
1
11 75°—1 * sJ
1 1 11 r1
AH(MW) Cpl t~ir»cr ioir»e/ n on /j " — IUJ IUI.UU** VJ.JJH
p /-v n —.^m/0 001 no 001 no
-©—
-©-86.403 -©~15.876
0-502—
1.184
94.5°
.; v..; v.293X)31
-^56° A.63 A.63 ^ r- x»o oc / no oc / no^JO CJU.uLu OD-4>UJ
^.fr-^ 1C O*7 C 1C Q"7C• •••••*p ou ib.o/b lo.o/D
—— 79° 0.502 0.502
52.56 O.U6 U1.9 0.258
G t»3 ^.«O<« U.U3/2-434
6-89 6.89 1.336 0.103
k n/O ic-ooc «c-oorrO /^ i^-^.^.^ i-v-z.z.-j 15.225
0-583 0-58382.987 Q2.987
i ———— R7° 9fl.^nz. n.?7R17.543 " 0-394 10-33
I „
= 123.5 MW = 401.6MW
Figure 5.22 Heat exchanger network for dehydrogenation process that can reach the energy targets.
151
temperature), and the rest of the stream is treated with hot
utility. The importance of this match is related mainly to
the hot utility, since the hot utility used on this stream is
at a very high level, that is, at a temperature over 710°C
(the higher the level the more expensive the utility is).
Therefore, this technique allows the hot utility on this
stream to be reduced from about 84 MW to almost 40 MW, i.e.
more than halved. This is achieved through the technique of
stream splitting carried out on stream No.I.
The design in Figure 5-22 involves an unwanted situation
shown by stream No.14. The supply and target temperatures of
this stream are 107 and 74 °C respectively, and this stream
is pinched at temperature of 106°C. Therefore a very small
part of this stream is above the pinch and the rest is below
it. Hence the part above the pinch can only be matched with
stream No.9 or stream No.4, since all other streams cannot be
matched with it due to the minimum driving force ( Tnin)
constraint. However, if it is matched with stream No. 9 or
stream No.4, it would result in an unnecessary match on both
the selected stream (No. 9 or No. 4) as well as stream No. 14
leading to an extra unit (Figure 5-22 exhibits the situation
when stream No.14 is matched with stream No.4)
If stream No.14 is not matched with any of these streams,
then cold utility above the pinch is needed, in order to
bring the temperature from 107 °C to 106°C. This would result
152
in a small violation in energy targets, and also a small
violation in minimum approach temperature between stream No.9
and stream No.l as a result of the cold utility usage above
the pinch. Therefore, to over come this complexity, stream
No. 14 will be totally shifted below the pinch. This shifting
is performed by increasing the minimum approach temperature
by 1 "C to 11 °C. This new approach temperature does not
significantly effect the energy targets, but produces a more
convenient heat exchanger network design as shown in Figure
5-23.
The design in Figure 5-23 has much improved the process
energy conservation, due to the maximum use of interprocess
heat recovery, for the given operating conditions.
However, heat recovery by interprocess heat exchange can
be improved even more, if possibilities for Process
improvements are found. This will be sought in the sections
that follow.
5.2.3 Process Improvements
The cold utility consumption in the design shown in
Figure 5-23 is considerably high (about 402 MW), and the main
reason for this, may be attributed to the demand of stream
No.2 (the crude styrene condenser). When the process
153
Stream No. cp =026 174.2°— ——— Pa 4 r 1 17,ID 565" —— (^ 1 x_p —
I2) Cp^C](5) rt(61
(7) (8) (U)
j
^(15) 520° ——— r (161 710°-— @ — -z-4 ———
r> X_£/(/>
172.76*
>* J w * * * ^ ^W CC / ^^
, 40.25 55A 101-6' wJ /
(9) 97° —————————(13) 110° ————— @ ———
(11)(12) 117°- ———— ® ———
0.583do) io7° —— e ——.^0 81>2 /
107°
7°129C 1245,•}ts~\f/-^ 1 /^N
-in
r i
^-9
-11X1(
1-77
I I105HI 157°—
'57°Ion0'OU
'l07°l »W '
1
1 o1 ?° °-+-
*-[96°
7 1 175°—
/ <*s
6°|
D6Jr
mAH(MW) Cp (^~
^ *-^\c 101 r\f t r\ ~^f\ /* * V^ %J 1 \J !• J J *4 ^J* ^^ ^ ™^
p /-\ /-s >9K _ m/o 101 no 001 no
4.63-e-86403
15^76-e—0-502-©- 1.184
,93.1°
J V.r 1[293.46 ^o* Jb **• bo 4>. bo
*^j/ oD-4Uo ob.4>Uo^ g"g^ ij- O"7C 1C Q*7e._„ ._^i-j^ ib.i5/b lb.o/b
—— 79° 0.502 0.502
52.56 O.U6 U1.9 0.258
N / r>O ^ / *% / /-> /-irN»^"~O ««3 ^..^Jt* U«U3/
24346.89 6.89
1.336 0-103\ T/O 4 1- -N-M- 41- O^f
O / L» IJ-ZZJ lvJ.^./.J
15-225 0-583 0-583
82-987 82.987(4) 100 "* O 1 v^ v_/ j/ ^o•ou^ \J-LIJ
17.6 ' 0.788 9.9
196'
=OCmm = ^02 MW
Figure 5.23 Heat exchanger network for dehydrogenation process with ATmin = 11°C.
154
composite curves are drawn as in Figure 5-24, stream No. 2
appears clearly to be the most dominant stream in terms of
cold utilities. This is because, this stream has a large heat
load and its supply and target temperatures (105 and 104°C
respectively) are not high enough to enable the stream to
exchange heat with other streams, nor for any other
integration proposals to be carried out. Therefore, this
stream needs to be utilized, in order to recover more energy.
5.2.3.1 Further Energy Recovery By Utilizing Stream No.2
Excluding the three streams No. 15, 16, and 4, all the
other seven cold streams involved in the network shown in
Figure 5-23 have a supply temperature less than 117 °C. Hence
raising the temperature of stream No.2 to 130 °C would enable
this stream to be matched with these other seven cold
streams. Thus, a further reduction will be gained in both hot
and cold utilities. This can be shown very clearly by
constructing the composite curves involving this temperature
increment as shown in Figure 5-25.
Figure 5-25 shows an increased overlapping between the hot
and cold composite curves, and this is a sign of increased
process to process heat exchange and of decreased hot and
cold utilities demand. For this modification the pinch point
location is changed, and this change has played a great part
155
01 en
650
o
1-50
0
§350
200 50
0
QH
min
*t23
.5M
r
QC m
in =
401
.6M
W
Stre
am n
o. 2
160
320
£80
Ent
halp
y (M
W)
6^0
Fig
ure
5.24
D
ehyd
roge
natio
n pr
oces
s co
mpo
site
cur
ves
befo
re
any
impr
ovem
ents
en
-a
650
o 7500
<D E o.
200 50
0
QC m
in =
302
MW
UO56
025
0 A2
0 E
ntha
lpy
(MW
)
Fig
ure
5.25
D
ehyd
roge
natr
on
proc
ess
com
posi
te
curv
es a
fte
r ra
isin
g th
ete
mpe
ratu
re
of
stre
am N
o. 2
to
13
0°C
.
in shifting almost all the cold streams below the pinch,
hence increasing the interprocess heat exchange.
This higher temperature for stream NO. 2 is achieved by
cooling stream No.l (crude styrene stream) to a temperature
of 130 "C instead of 105 °C, and increasing this streams
pressure to raise the condensing temperature of the gas to
130 °C instead of the old condensing temperature of 105 °C.
The resultant energy targets and pinch location from this
modification are as follows;
The hot utility is 40 MW (an 82% reduction on the original
design), and the cold utility is 302 MW (a 40% reduction on
the original design), and the pinch location is at interval
temperature of 559.5 °C.
These energy targets and pinch location have resulted from
deriving the problem table of the process for a minimum
approach temperature of 11 °C. But in this situation and
after the above changes have been made on the process, it is
possible to bring back the minimum approach temperature to 10
°C (the original assumed temperature), now the pinch location
will be at interval temperature of 560 °C. Table 5-5 shows
the problem table that is derived for minimum approach
temperature of 10 °C.
158
Interval Temperature (°C)
715
560
525
165
125
124
122
121
115
112
111
102
101
80
79
75
74
69
62
54
53
52
51
Heat Flow (MW)
40.0
0.0
4.9
3.2
8.2
319.9
319.4
318.5
316.9
315.8
232.4
229.0
221.9
216.9
201.4
200.1
200.2
198.6
195.9
195.2
195.2
281.6
302.1
Table 5-5, The problem table for the dehydrogenation process after cooling stream No. 1 to 130 C, derived for minimum approach temperature of 10
159
The heat exchanger network design for minimum approach
temperature of 10 " C that maintains the energy targets in
Table 5-5 is shown in Figure 5-26.
Not only does this design maximize heat recovery, it also
minimize the number of units that are used to maintain the
energy targets. The minimum number can be calculated from the
following equation;
Urn in = N - 1 ...................................... (6-1)
Umin = 18 - 1 = 17 units
This is the number of units shown in Figure 5-26. The
reason for this is that, this design offers an opportunity
for the tick off rule to be completely applied, as explained
in the following section.
5.2.3.1.1 Tick Off Rule Application
The first match below the pinch in Figure 5-26 is between
stream No.1 and 15. The load on this match is maximized until
stream No.15 is satisfied, thus ticked off. The second match
is between stream No. 1 and 16, this match ticks off stream
No.16. Similarly the third match is between stream No.1 and
4, this match eliminates stream No.l.
160
565°
1 Stream NO. |C Ph*o.u~ 189
( 1 ) 565*(2) (5) (6) (7) 18) (14)
(15)(16) —— ® —— — (3) 39 '99
(9)(13)(11) (12)
(10) ,,,
5 /T*CpaOJ^130*-*• •«• O57 -
57°—*J /
107°-
52 O2**jt. \j—— n
s? — v
•5C
173° > f\ ^\^\^\
-^
e lJ—
-©--
IN,
vsn;-fc-1.23
ratn 101.974°- ——
97° ———
110° ———
?^
107° ——— i
s\.^
^-^-1 1.336
-^--1 ^15.225
0.583
MW AH(MW) Cp Uc J
J — IJU !./*» U-*»
,/v.
k_-^^n6.89
88-87* •\ t /
-^v.j^r~\£.'d Jiz 01^ N.78— *• 56 4-63 4-63
_ 1—^.0 nr / «o 0^ / /^*i— *• 57 86-403 OD-403
^56 Ib^i/b Ib-o/b— *-7Q° n ^n? nsn?
/ 3 U-«?U^ U-JVJ^.
— 7/° 1 9T n m?/»» l-^o U-Uo/
—— 160° 52-56 0-146—— 160° 141-9 0-258A / /~P i/o/ r\ r\ r\*^~O ««y ^.^j** u-ua/
yb b'oy b>oy———— 07° 1 o*3c n in^s/ I.JJD u-iuj
11 0 ir OOC 1 C OOC/4 Ib'ZZb lb-^^bo . .11 c n c 0*5 n c o *5lib 0-L)o3 Do 83
——— 1(^C 0*1 00*7 o-"> no*?lUb o2-9o7 o2-987 -i ——— c;70 ofl.Tn/. n.07c:
555°
OH = QHmin = 39.9 MW = QCmin = 302.4 MW
Figure 5.26 Heat exchanger network for maximum energy recovery when crude styrene stream is cooled to 130° and AT mm is taken to be 10°C.
161
All remaining cold streams are ticked off by matching them
with stream No. 2, since this stream has enough heat load to
satisfy all these remaining cold streams. After all cold
streams are satisfied, stream No. 2 will be ticked off by
using a cold utility, as are the other hot streams.
Above the pinch there is one stream only (stream No.16).
This stream is a cold stream, and it is satisfied (ticked
off) by using a hot utility, since there is no hot stream at
a high enough temperature to match it with.
Therefore, the network design in Figure 5-26 represents
one of the most convenient designs that can be derived for
this process. Since it uses considerably less energy than the
original network design, this is achieved using the minimum
number of units possible, as well as satisfying most hot and
cold stream demands using only one unit which would make the
design easy to operate and control.
However the cold utility consumed in this design still
high, and this will be examined in the following section to
find out whether further reduction in cold utility is
possible.
162
5.2.3.2 The Reduction of Utility Consumption By Utility
Generation
Almost all the cold streams have taken their heat demand
from other hot streams in the process by interprocess heat
exchange. This interprocess heat exchange has reduced the hot
utilities considerably to about 40 MW. Due to this process to
process heat exchange the cold utility is decreased too, but
there is still a considerable demand of 300 MW. The reason
for this high consumption of cold utility is that, the hot
streams have much more heat than the cold streams require,
therefore the unrecovered heat creates a high cold utility
demand.
The suggested technique to reduce this high consumption of
cold utility is, to recover the available heat in a waste
heat boiler, in which low pressure steam can be produced at
120 °C and 2 bar. In general the recovery of heat as steam is
very attractive because, this steam can be:
1- Used as a hot utility where ever it is needed in the
process, thus reducing the hot and cold utilities at the
same time.
2- Exported to another process.
3- Passed into turbines to generate power.
The generation of steam in this section is looked at from
163
the angle that it has great impact on the reduction of cold
utility. The interaction of this steam within the process
will be examined in the next chapter.
Because the minimum approach temperature is 10 °C, then
the highest temperature for the generation of steam is 120 °C
Since most of the heat available in the process is at a
temperature of 130 °C in stream No.2.
The right amount of steam raising must be calculated in
order to maintain an operable heat exchanger network design.
5.2.3.2.1 Cold Utility Reduction by Steam Generation
The maximum amount of steam raising will be estimated by
the aid of the process grand composite curve. The process
grand composite curve before steam raising is shown as a
solid line in Figure 5-27.
The actual low pressure steam temperature has to be
converted into an interval temperature. As the steam raised
is a cold stream (accepting heat), then the interval
temperature is 125 °C (120 + 10/2). At 125 °C the grand
composite curve indicates that the maximum amount of heat
available for steam raising is 195 MW. The effect of the
steam raising on the grand composite curve at a temperature
below 125 °C is shown as a dotted line in Figure 5-27.
164
a> CJl
800
700
o a J/) £5
00
J400
o l_
0>
C
L £300
200
100
o
Qcm
in =
107
MW
QC
min
= 3
02 M
W
0
Fig
ure
5.
27
50
120
160
200
240
Ent
halp
y (M
W)
280
320
360
The
dehy
drog
enat
ion
proc
ess
gran
d co
mpo
site
cu
rve
befo
re
and
afte
r ra
isin
g th
e m
axim
um
amou
nt
of
stea
m.
The modified grand composite curve reveals that the heat
demand on cooling water is reduced by 195 MW to 107 MW. This
maximization of steam raising capability causes a second
pinch point (utility pinch) to occur at an interval
temperature of 54 °C (shown on the modified grand composite
curve). The effect of maximizing steam generation and
creating another pinch point upon the heat exchanger network
design is shown in Figure 5-28.
Although the design in Figure 5-28 generates the maximum
amount of steam, it displays a situation where the process is
difficult to control for the following reasons;
1- Stream No.8 (benzene column condenser) is cooled by
matching it with stream No. 3 (the feed to benzene-toluene
column). This is not advisable, since the reboiler for the
benzene column (stream No.12) is already integrated with
the process, therefore, it is not recommended to integrate
the condenser of this column as well (for operability
reason). This saves creating an extra unit on stream No.3.
2- Stream No.14 (the feed to the styrene finishing column) is
cooled by matching it with stream No.4. Stream No. 4
represents the ethylbenzene stream that is sent to the
dehydrogenator, which contains the ethylbenzene produced
by the alkylation process plus that recycled from the
dehydrogenation process. Thus, it may not be advisable to
166
565 59<
Stream No. 189.5°
49^
MW AH(MW) Cp( t
M] nX— P-l%036
(2 ) ' 130°" (5) | (6) Im [(14) 1 107-
1 1
(15) 1520^71 Q° 1
X J_o «"»rioi173° \s\ r\ /-\r\
^~**152-36 s
m 39-99 1 .vo 101.9 1
J \
\IJ i i /A — O 1 o 1-93
JTT.
^(9) |97°- ———— (,-| (13) 1110° ———— H^
J\
>i
I _n 1-336 fill '75^-1 ' 1 1 1
MO 1 11 7° / \\l] ill/— —————————— t/l
J
~\
1 0-583(10) 11072. ——————— -
(4 ) il60- ————— O w
__.
^^
4t
7\:
^
_..._ — „_ (OU .
^^/^. <onP.J \.
-e —-P-
S-^'n0.502
15.22
61>1 o 19-5A 7.534 1 PI 'ion ^ _
.5° -^-i
3C
•< n 82-9•v
1.23 f
r IC7° /^N e-/r°3/ l^J ->D
crqO ??3 T7°A-x 3/
C7° r^ ^f^b/- - vj>- "— bb 15^76•^ *"» O-"79°
lou
AQ96°
97°
116°106°
57°
!i_11Q0I Low pressure steam 195.51 0 The maximum
555 amount of steam raising
174312
* 4.63
86.4033 15.876
0.5021.23
52.56141.92-4346-89
1.33615-2250.583
82.98728.304
195.2
0-43124.6386.40315.876
0.5021-23
0.1460.2580-0976.890.10315.2250.58382-9870.275
195-2
O H = 39.99=*40 MW Q C = QCmin = 106-909 MW
Figure 5.28 Heat exchanger network design for dehydrogenationprocess when maximum amount of steam is raised.
167
connect this stream to stream No. 14, due to the
considerable pipe work involved, considering that the heat
recovered is only 1.23 MW which is not very high. In
addition this match requires an extra unit on stream No.4.
Therefore, to produce a better design, the amount of steam
raised has to be reduced. The reduction of steam generation
will remove the utility pinch, thus the constraints on the
process will be lessened.
Inspection of the designs in Figures 5-28 and 5-26 shows
that the main difference between them beside raising steam
is, that for the maximum steam raising case (Figure 5-28)
streams No.8 and 14 are matched with process streams.
Therefore, if these streams are rematched with cold utility,
the steam generation will be reduced to 193 MW. Rematching
streams No. 8 and 14 to cold utility does not increase the
cold utility much, since the load on these two streams is
about 1.7 MW which is not great. This case will produce a
design similar to the one in Figure 5-26, but with a
possibility of generating steam equivalent to 193 MW without
creating a utility pinch. This design is shown in Figure 5-
29.
The energy recovery at this stage is effectively at its
practical maximum, since no more improvements can be carried
on the process for further energy integration. The improved
168
565'
Stream No. |Cph =o.u tfl(1) 565° — f^
(2) 130 I(5) 57° —\ j i i • f\1 C \ C Q(6) 58
(7) "57°1 / i i -> ' f Q i i on°(o J I ouM/ \ ho/? 0 —\ |t» I ' 1 \J /
1
(15) |520°—
^
p — v95° VlZ?*85"
173* ^T^^
s
>2-5( s(16) —— @ —— i ——— — i o 1 39-99 1 7/0 J^?_,\ J ; . /*» tQl i 97°- ——V 3 / | J f
(13) '110° ———
(11) 75°-* ——III; I
(12). |117°- —— (10) 1107° ——— b —— -
./v.
-^^-1
/ •©^
8|4
15-83
o-so:1.23
^0-583^
1 0 82-487 (4) '160°"* — 6 w
./
J V.
-i —— I
J V.J V.J V.
^ .-(^1243
•*> ———— 1
649s1336
N** —————— 1 15-225
, 88-8° J
J V.
^. .
I « 8 -? 64
130*
160
96°
97°
116°i f\ x* O
57°
555C Low pressure steam 193.768
OH = QHmin = 39.99MW
AH(MW)
174312
£.63
86-40315.8760.5021.23
52.723U2.1022.4346.89
1.33615-2250.58382.98728.304193-4
0.4312
4.6386.40315.8760.5020.037
0.1460.2580.0976.89
0.10315.2250.583
82.9870.275193-4
=108.6MW
Figure 5. 29 Heat exchanger network for maximum energyrecovery with steam raising.
169
dehydrogenation process flowsheet that represents this heat
exchanger network is shown in Figure 5-30. This flowsheet
exhibits all the improvements that been carried out on the
process to produce an energy integrated dehydrogenation
process.
5.2.4 Dehydrogenation Process Utility Levels
The modified dehydrogenation process grand composite curve
is shown in Figure 5-31. This curve reveals that the process
consumes about 40 and 108 MW as a hot and cold utilities
respectively.
The hot utility has to be introduced to the process at a
considerably high temperature level (over 720°C), Therefore,
a furnace has to be used to satisfy the heat demand of the
process at such a temperature. This requirement will be
considered in more detail in the next chapter, when the
interaction between process and utility is sought.
The cold utility is to be introduced to the process at a
temperature level below 45 °C. The temperature levels of the
hot and cold utilities are estimated from the modified grand
composite curve.
170
f-^" Stre
£fr 2
Vj-"0u*0
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485151
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"76629.2
76629.2
Furnace Dehydrogenator Gravity Separator Benzene and Toluene —— Column
Ethylbenzene Column
Styrene Column
Benzene Column
R__Ethylbenzene
LR steam raising
Stream locationHeatcontentlMWJwith reference temperature of 25°&pressure of I bar
Tempera tu re (*C)
A
5555
b65
A,
189.4
565
A?
366.14
565
A3
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189.5
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A 7
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239.6
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A9
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130
Aio
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An
43.71
130*•' mm
Ai?
2-973
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An
8.^36
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Au
1.647
130
Ais
101.5
130
AIR
18.62
130
Ai?
0-723
130
AM
0-14
130
Aiq
3-402
130
A?n
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130
ATI
0-311
130
A??
1.546
130
A 73
0543
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A24
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A25
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B
67.73
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Bl
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C
326£
160
Ci
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555
C2
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D
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49
E
0.725
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F
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F1
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C
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G1
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56
G2
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H
9954
58
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11.3
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H3
3.071
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H4
9.695
120
H 5
29-21
138
H6
31.33
160
I
18-02
57
H
1.073
56
12
1-073
56
J
0-625
80
Jl
0.031
79
J2
0.092
79
K
6-879
97
Kl
8.216
110
L
2.96
107
Li
1-728
74
M
0-005
75
N
0-098
117
P
362-9
160
R
1-233
57
z-3o Dehydrogenation process flowsheet after being energy integrated
ZLl
Tl ia'c
01
00
Temperature Interval (°C)K>
oSf\
CO
3 5"
0OD
'oi3 5' it
\
3T3
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oD
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nc
5.2.5 Concluding Remarks
Process energy integration on the dehydrogenation process
has resulted in a great reduction in utility consumption. The
process utility requirements before energy integration were
500.2 MW as cold utility and 221.6 MW as hot utility. After
energy integration, the utility requirements are 108.6 MW as
cold utility (78.3% less than the original usage), and 40 MW
as hot utility (81.9% less than the original usage).
173
CHAPTER SIX
The Utility Interface With the Process Design
The results of process design always give the demand for
utilities, and the utility system is then designed with
respect to the process. Process energy integration can be
introduced at this stage to seek the design that requires the
minimum energy, and the corresponding levels at which the
utility should enter the process can also be determined.
The work in this chapter will deal with the design of the
process and utility simultaneously. The objective is to
assign the most efficient way for the utility to be
introduced to the process. Consequently the heat exchanger
network design should become more effective, resulting in a
reduced demand for utility imports.
6.1 Dehydrogenation Process and Utility Consumption
In the previous chapter the dehydrogenation process was
integrated, and the network design that is compatible with
its minimum energy requirement is shown in Figure 5-29. The
hot utility required for this network amounts to 40 MW. This
hot utility needs to enter the process at a very high
temperature level (over 710 °C). Therefore, the utility is
supplied by flue gas (obtained from fuel combustion in a
174
furnace), since the utility supply by other techniques such
as, turbine exhaust and/or direct steam heating can not be
used.
The gases that are vented from the condensation section in
the dehydrogenation process itself can be used as a fuel in
the furnace unit< 77 - 78 >. Hence the flame temperature of this
fuel needs to be calculated (since it is very important when
the flue gas is introduced as a process stream, as will be
seen later). The calculation of the flame temperature is
shown in appendix (E), and this calculation shows that this
gas has a flame temperature of 1500 °C.
6.2 The Use of Hot Utility and Stack Loss
The region above the pinch for the energy integrated
dehydrogenation section (Figure 5-29) along with the hot
utility system is shown in Figure 6-1. The utility needed is
40 MW, and it is supplied by flue gas available at 1500 °C.
The highest recovery of heat from the flue gas should take
the flue gas to the lowest allowable temperature which is 160
"C (the acid dew point limit).
In Figure 6-1, the flue gas leaves the furnace at 565 °C,
and thus there is a high wastage of heat through the stack.
The heat lost in the stack and the amount of fuel consumed
175
565'
II
The area above the pinch The area below
the pinch
710*
555
710
555°
555°
T stack -- 565°C
Figure 6.1 The network design in the area above the pinch including hot utility.
176
are determined as in the following procedure.
The heat recovered from the flue gas in the network is 40
MW. This recovery of heat has reduced the flue gas
temperature from 1500° to 565 "C, and thus the value of CP is
found to be
Cp = 40 / (1500 - 565)
CP = 0.043 MW/°C
Then Q = 0.043 (1500 - 160) = 57.62 MW
therefore the heat lost in the stack = 57.62 - 40 = 17.62 MW
To estimate the amount of fuel consumed in the furnace,
the flue gas has to be taken to a temperature of 25 °C
(ambient temperature) in order to calculate the total heat
content in the flue gas.
Therefore Q = 0.043 (1500 - 25) = 63.425 MW
assuming 5% of the total heat is lost through the walls would
result in Q = 66.6 MW .
The net calorific value of the fuel used is 14.56 MJ/m3 ,
thus the amount of fuel consumed will be
66.6Fuel consumed = —————— = 4.57 m3 /s
14,56
In general different techniques can be applied to reduce
the loss of heat in the stack. Of these techniques, the use
177
of a waste heat boiler for steam raising, or preheating the
air required for combustion before entering the furnace are
most commonly used. The steam raising scheme may be
considered depending on the need of the process for steam.
Preheating the air would reduce the fuel consumption in the
furnace, thus this scheme is considered when the amount of
fuel consumption matters.
However these two schemes will not be adopted, because for
the first scheme the amount of steam that has to be generated
by recovering the heat in the stack is not significant, and
for the second scheme the fuel consumption in the furnace is
not the main concern, since it is produced to excess in the
process as a by product, and this amounts to 9.52 m3 /s (0.425
kmol/s). Also these two schemes do not further improve the
process heat exchanger network, and do not further reduce the
utility consumed by the process over the 17.6 MW available in
the stack gas. Moreover, the above approaches do not have any
effect on the boiler house in terms of reducing the fuel
consumed in the boiler house or shutting down one or all the
boilers included in the boiler house. Notice that the fuel
released by the process has a small calorific value, thus
using it in the boiler house will not reduce the import of
fuel (for the boiler house) significantly. Therefore, the
approach that is adapted in this chapter will contribute to
all the criteria mentioned, in addition to the minimization
of heat lost in the stack. Hence the furnace need not be
178
taken as an after thought (as the techniques described above
do), but needs to be considered as an integral part in the
context of overall process design.
Such an approach will deal with the flue gas as a normal
hot process stream to be represented as a grid with the rest
of the process streams. This approach would usually result in
a reduction in fuel consumption* 7 9 > . But the fuel is
available in the process as a by-product, hence in this
approach it is tackled in such a way that the fuel
consumption in the furnace is increased, in order to help
maximize the interprocess heat exchange and the other
criteria mentioned above. The network design that involves
the flue gas as a hot process stream is shown in Figure 6-2.
6.3 Flue Gas in Process Context
As mentioned above the lowest temperature that the flue
gas can be taken to is 160 °C. Therefore the supply and
target temperatures of the flue gas stream, as shown in
Figure 6-2, are respectively 1500° and 160 "C. The heat
capacity flowrate of the flue gas stream is not known, and
depends on the load that the stream needs to be matched with.
The network in Figure 6-2 is designed so that the heat
demand of stream No.16 is totally satisfied by matching it
179
1500C
Stream No. IFG
(1)(2)(5)(6)(7)(8)(14)
(15)
(16) (3)(9) (13)
(11! (12) (10!
U) LR
—————— trrr0 /""N^33.6'JU'J V^
130°-^.— »O57 —CC°T .bo57° —«^ /
O U
107°-
— .* — .O52 0-*- 71 0°-*•71 °74 -*-97°*
110°-*-75°—
117°—1fY7°.«f
./
^
y
1
^52-56
"^ ——— 1
J
AJ6386-4
5*7
)602
1-23
^n
——————————————— , ou
r\/^^i^>
^i^UI.9
2-43 4 /-1
j\.
O —————— 6'89G 1
^ V.
s1.336.
yv.
^1^225^
j
O N 0-583
IU/ ^\J Oj 82.987 J
V — IOU
.,,-^T^ «nn°
^IDU -• O 28.3 J
7 — 1^. J
r- *»O
—— 57°*•/ /
» rr°* JO
_ — o—— *-79
——— 160°
n°—— 96°——— 97 C
-7/0
106°
57°
"x A « j-«.O
490
Low pressure steam 202-5
QC = 108-6 MW
AH(MW)MW
141.9 CP = ?
1.74 0.4312 3124.63 4.63
86-403 86-40315-876 15.8760-502 0-5021.23 0.037
52.56141.92-4346.891.33615.2250.58382-98728.304
202.5
0.1460.2580.0976.890.10315.225
0.58382.9870.275
202.5
Figure 6. 2 The network design of dehydrogenation processafter introducing the flue gas as a process stream.
180
with the flue gas stream (this has resulted in a stack
temperature of 170 °C). Which means that the steam stream
(stream NO.16) is heated up directly by the furnace to reach
the temperature of 710 °C. This supercedes the approach in
the previous chapter which is shown in Figure 5-29, in which
stream No. 16 takes some of its heat demand by heat exchange
with crude styrene stream (stream NO.1), and is then sent to
the furnace for further heating. This part of the process
before and after heating stream No.16 directly by the furnace
is illustrated in Figures 6-3a and b respectively.
The amount of the fuel required in the furnace in the
approach shown in Figure 6-2 is 11.35 m3 /s, as shown in the
following calculations;
The heat load that has to be given from the flue gas
stream to stream No.16 is 141.9 MW (the heat demand of stream
No.16).
141.9Therefore CP = ———————————— = 0.1067 MW/°C
1500 - 170
thus the total heat content in the flue gas will be
Q = 0.1067 (1500 - 25) = 157.38 MW
assuming 5% loss through the walls would result in a net heat
content of 165.25 MW.
165.25Therefore the fuel consumed = ———————— = 11.35 m3 /s
14.56
This amount of fuel is more than the fuel available in the
181
Steam from the boiler house at 160* and
Stack at 565°
Figure 6.3a Stream No. 16 (steam stream) before direct superheat in the furnace.
Steam at 160 EB (4VI I IUU
It
90*/.|1 £
(16)1I1
Stack at VXs^s
,'c ->^
*"->s''
s
Furnace
70°
Tjg°l^^^iiiiii
IU /•
/^~ ~"^\
O-1—DCQJ
OTD
£. QiQ
(1
"
565° SL-±L-Qij
'
1
(4)
160°
I 433.6° f
520°
Figure 6.3b Steam No. 16 (steam stream) after direct heating in the furnace.
182
process by 1.83 m3 /s. Therefore, some natural gas (which is
used in the boiler house for the styrene plant) may be mixed
with the available fuel to cover the amount required. The
amount of natural gas that is required is not 1.83 m3 /s,
because the two fuels differ in their net calorific values.
Therefore, the exact amount of natural gas required is
calculated as follow;
The total heat required in the furnace is 165.25 MW, but
the total heat that can be supplied by the available fuel is
9.52 x 14.56 = 138.6 MW. Therefore the natural gas is
required to supply the difference which 26.65 MW. The net
calorific value of the natural gas is 37.67 MJ/m3 , hence
dividing the heat required by the calorific value would
result the amount of natural gas required which is 0.7 m3 /s
(0.03 kmol/s).
Analyzing the network shown in Figure 6-2 shows that the
increased fuel consumption in the furnace (which has resulted
in an import of only 0.7 m3 /s) has improved the overall heat
recovery in the network design, and has given it many
advantages, as will be discussed in the following section.
6.3.1 The Network Design Analysis After Introducing the Flue Gas As a Process Stream
The network design shown in Figure 6-2 (after introducing
183
the flue gas as a process stream) is compared with the
network design shown in Figure 5-29 (before utility
interaction with the process is considered). The comparison
is taken in order to measure the network improvements due to
introducing the utility in a different form. The advantages
that the network in Figure 6-2 has over the one in Figure 5-
29 are;
1- The flue gas stream is cooled to a temperature of 170 °C,
whereas the old maintained temperature is 565 °C. Which
means the energy lost in the stack is reduced to nearly
the minimum.
2- In Figure 5-29 stream No.16 takes about 102 MW from
stream No. 1 by heat exchange. Therefore, by heating up
stream No.16 totally by the flue gas, releases about 102
MW in stream No. 1 at a ^asonably high temperature level
(between 433.6° and 178.6 °C), and this load may be
utilized elsewhere. Wherever this heat load is utilized,
more energy saving and a better network design will
result. The utilization of the heat load on stream No.l
will be considered in the next section.
3- Stream No.4 in Figure 5-29 takes some of its heat
demand from stream No.l and the rest is satisfied by
matching it with stream No. 2. This is so because the heat
content in stream No.l was not enough to satisfy stream
184
No. 4 totally, due to the heat that had to be given to
stream No.16. After matching stream No.16 with the flue
gas, it is possible to satisfy stream No.4 totally by heat
exchange with stream No.l. Therefore, the heat demand of
stream No. 4 is provided by using one not two units, and
this reduction in the number of units reduces the capital
cost and gives a stream that is much easier to control.
4- Because stream No.4 has taken its heat demand totally from
stream No.l, some heat (9 MW) is available in stream No. 2
and is used to maximize the low pressure steam generation.
Therefore, the amount of low pressure steam generated in
the process is increased from 193.7 MW (the old maintained
amount) to 202,5 MW.
The last two points (3 and 4) can be considered as a part
of the utilization of the heat content that is left in stream
No.l (due matching stream No. 16 with the flue gas). This is
so because an extra 9 MW is taken from this heat content to
fulfil the requirement of stream No.4. Thus resulting in a
one unit match between stream No. 4 and 1 that is capable of
totally satisfying stream No.4.
In the old network shown in Figure 5-29 stream No. 4 is
matched with streams No.l and 2, and because the new approach
shown in Figure 6-2 has enabled stream No. 4 to be satisfied
by matching it with only stream No.l, then about 9 MW is left
185
unused in stream No. 2. This 9 MW is then used for low
pressure steam raising.
6.3.1.1 The Utilization of the Heat Load on Stream No.l
As mentioned above the amount of heat content that is
available in stream No.l is 102 MW. Some 9 MW is utilized
with stream No. 4, therefore, the amount of heat left is 93
MW. This heat content is shown on stream No.l in Figure 6-2
as a cross sign in the middle of a circle, and is between
temperature levels of 433.6° and 201 °C.
Further analysis of the network shown in Figure 6-2 (the
network design of dehydrogenation process) reveals that there
is no scope for utilizing the remaining heat content in
stream No.l by heat exchange with any other cold stream in
the dehydrogenation process. Since all the heat demands of
all cold streams have been satisfied by interprocess heat
exchange. Therefore, the alkylation process will be
considered for further energy integration, as there is a
considerable heat content in the dehydrogenation process that
may be exported to the alkylation process.
The energy integration of alkylation process (in chapter
5) left a demand for 14.4 MW of hot utility. The hot utility
is mainly needed for the reboilers associated with the second
186
and third columns. Therefore, by matching stream No.l (crude
styrene stream) to these two reboilers, the heat content
would be reduced to 78.6 MW. Due to this heat exchange the
temperature of stream No.l is cooled from 433.6° to 397.6 "C.
Stream No.l still needs to be cooled to a temperature of 201
°C and has a heat content of 78.6 MW to be utilized.
At this stage neither the alkylation process nor the
dehydrogenation process can accept any more heat, since all
the heat demands of their cold streams have been satisfied by
interprocess heat exchange. Therefore, the heat load of 78.6
MW is recovered in a waste heat boiler in which intermediate
pressure steam at 160 °C and 6 bar is raised. The utilization
of the heat load on stream No.l is illustrated in Figure 6-4.
The process already raises 202 MW of low pressure steam,
hence the utilization of this steam together with the amount
generated at an intermediate pressure will be examined in the
next section.
6.4 The Utilization of the Steam Raised in the Process
Due to the integration techniques carried out on the
styrene process (in this and the previous chapters), the only
unit in the total styrene process that requires steam is the
dehydrogenator.
187
Ste
am a
t 16
0°10
V.
160C
(16)
1S
tack
at
170
°
Al_
_
00 00Fu
rnac
e
710°
>
Dehydrogenator
.._ ———————— A
^
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555
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433.
6
Sepa
ratio
n sy
stem
at
alk
ylat
ion
proc
ess
160°
57°
EB
78.6
MW
of
med
ium
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essu
re s
team
rai
sing
397.
6°To
con
dens
atio
n se
ctio
n
520°
Fig
ure
6.Th
e util
izatio
n
of
Stre
am
No.
1
.
The mass of steam needed is 15 times the mass of
ethylbenzene fed to the reactor, and this amounts to 362 MW
(in terms of energy). The steam is supplied from the main
boiler at 160 °C and 6 bar. The steam supplied is normally
divided into two branches before it enters the
dehydrogenator. The branches carry 10 and 90% of the total
amount of steam. The branch with 10% of the steam flow is
mixed with the ethylbenzene stream and is then heated up by
heat exchange with the crude styrene stream before entering
the dehydrogenator. The branch with 90% of the steam flow is
superheated in a furnace, and then sent to the
dehydrogenator.
The amount of steam raised in the process is at two levels
and these are, 78.6 MW at 160 °C and 202 MW at 120 °C. In
order to make use of these amounts of steam, the steam that
is imported from the main boiler will be cut down, so that
the steam produced in the process may be used instead. This
technique will have a significant reduction in the utility
imported from the boiler house, in addition to the
availability of the utility at no fuel cost.
The medium pressure steam that is produced by the process,
which is at a level of 160 °C, can be used directly, since it
is at the same level as the steam that would otherwise be
imported from the boiler house. However, the steam that is at
a level of 120 °C must pass through a compressor to raise its
189
temperature and pressure, in order for it to be at the same
level as the steam imported, so that it can replace it.
The power input to the compressor is equal to the
difference between the heat contents of the output and the
input to the compressor. The input is saturated steam at a
temperature of 120 °C, and the output is saturated steam at a
temperature of 160 °C, The values of the heat content are
taken from steam tables* 80 >. Calculations show that the power
input required to drive the compressor, with the assumption
of 90% efficiency, is 5 MW.
From the optimization point of view, the 5 MW used in the
compressor will not be considered as a great cost, because it
has shifted the 202 MW of low pressure steam from a
temperature of 120 °C to a temperature level of 160 °C. The
transfer of the level of the steam from 120° to 160 °C would
result in a small reduction in the heat content, due to the
small difference in the latent heat of vaporization, and this
is also considered to be negligible.
Therefore, the total amount of steam available at 160 °C
in the process is 280.6 MW (the sum of the amounts of steam
raised in the process). The total amount of steam required in
the dehydrogenator is 362 MW, Therefore by using the steam
generated in the process, the need for steam from the boiler
house will be reduced to 81.4 MW, and this is a significant
190
reduction. Figure 6-5 illustrate the steam distribution in
the dehydrogenation process as a consequence of steam raising
within the process.
The effect of energy integration within the styrene
process using a system of combined heat and power will be
undertaken in the next chapter.
6.5 Concluding Remarks
The utility interaction with process design has led to a
considerable reduction in utility consumption. This reduction
is achieved by introducing the flue gas as a process stream.
The amount of fuel consumption in the furnace is increased
(in which the total by product fuel available in the process
is used, in addition to small portion of fuel that is
imported) to handle the extra load on the furnace, due to
superheating the steam stream directly in the furnace. This
step has decreased the loss of heat in the stack to nearly
the minimum, as well as resulting in releasing considerable
heat load in the process which is utilized in different ways,
as presented in the text, maintaining a higher recovery of
heat.
Due to this interaction the amount of steam raised in the
process is increased, and by utilizing this steam in the
191
(O
ro
Alk
yla
tion
se
para
tion
Ste
am i
mpo
rt
from
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ess
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process, the need for importing steam is reduced by 77.5%.
Therefore, the increment of fuel consumption in the furnace
is justified by the great reduction of fuel import for the
boiler house, as will be seen in the next chapter.
193
CHAPTER SEVEN
The Effect of Process Integration on the Steam and Power
System
The steam supplied to a process is usually produced in the
main boiler house. Before the steam reaches the process, it
passes through a central turbine in order to give some credit
for power generation. The exhaust from this turbine is
usually at different levels, in order to meet the process
requirements.
If the process is energy integrated, then the amount of
steam needed will be reduced, in addition the level at which
the steam enters the process may be lower. Consequently, the
way that the process is connected to the utility system will
be different, as will be both the scope for power generation
and the utility fuel consumption. Therefore, the work in this
chapter has taken the system of combined heat and power
generation in the overall site context to determine how much
this is affected by the overall efficiency of energy supply
and demand.
7.1 Combined Heat and Power System in the Styrene Process
A combined heat and power system supporting the styrene
process, as any other process, is designed to satisfy the
194
process need for steam as well as some of the site power
demand. Thus, the efficiency of this system (in terms of fuel
consumption and power generation) will depend a great deal on
the process requirements for steam, and the level at which
this steam is introduced into the process.
Therefore, the heat and power system in the styrene plant
will be examined before and after energy integration, in
order to assign the benefits that the system has due to
energy integration being applied to the process design.
7.1.1 Application of a Combined Heat and Power System in the
Styrene Plant Before Energy Integration
The styrene plant is divided into two main processes,
these are the alkylation and dehydrogenation processes. These
two processes have different demands for steam in both
amounts consumed and pressure levels.
in the alkylation process the main steam requirements is
to satisfy the need of the reboilers in its distillation
train. The amount of steam required is 66.1 MW, and is at a
temperature and pressure of 235 °C and 30 bar respectively.
Thus, only one level of steam is needed in the alkylation
process and it is required at a high pressure. As the latent
heat of vaporization for steam at 235 °C and 30 bar is 32
195
MJ/mol, then the amount of steam required in terms of moles
is equal to 2.0 kmol/s.
The dehydrogenation process requires steam at two levels,
intermediate and low pressures. The intermediate pressure
steam is needed in the dehydrogenator as a diluent, whereas
the low pressure steam is needed for the reboilers in the
distillation train. The amount of intermediate pressure steam
required is 362 MW, and is at temperature and pressure of 160
°C and 6 bar respectively. The amount of low pressure steam
required is 106 MW, and is at temperature and pressure of 125
°C and 2.3 bar respectively.
The latent heat of vaporization at 160 °C is equal to 37.5
MJ/kmol, therefore the amount of intermediate pressure steam
in terms of moles is 9.6 kmol/s. The latent heat of
vaporization at 125 °C is equal to 39.5 MJ/kmol, thus the
amount of low pressure steam in terms of moles is 2.7 kmol/s.
From the calculations above, the total amount of steam
required for the styrene plant is 14.3 kmol/s. This steam is
produced at the main boiler house, which consists of two
boilers. Each of these boilers produces 7.15 kmol/s. The two
boilers produce steam at a temperature and pressure of 450 °C
and 90 bar. The steam is fed to the central turbine in the
power station in order to generate shaft work. The steam
extracted from the power station is at three levels (HP, MP
196
and LP) , and are distributed around the plant. Figure 7-1
illustrates the connection between the process and the system
for combined heat and power.
7.1.1.1 Power Generation in the Plant
The amount of power generated in the plant depends on the
heat input to the central turbine and the heat output. These
values are estimated with the aid of steam tables. Table 7-1
shows the enthalpy values for the steam entering and leaving
the central turbine at different pressure levels.
Therefore, the total heat input is 838 MW, and the total
heat output is 710.4 MW.
For a turbine
W = heat input - heat output
Therefore W = 838 - 710.4 = 127.6 MW
Taking into account that only 80% of this energy is
transferred into electricity via the turbo generator, the
total electricity available in the plant is 102.1 MW.
7.1.1.2 Fuel Consumption in the Boiler house
The amount of fuel consumption in the boiler house is one
197
Fuel!U tr^3
s
Fuel
11.4 2lls
Boiler (1)
Boiler (2)
kmol 7.15 ~i~
(
- 1R kmol 7.15^
——— ». ———
i
I Stand by I Boiler
o O
O
W= 102.08MW
o <v
Figure 7.1 The boiler house and power station in connection with styrene process before energy integratl'on.
198
Steam Level
Very High
Pressure (VHP)
High Pressure
(HP)
Medium Pressure
(MP)
Low Pressure
(LP)
Temperature (°C)
450
235
160
125
Pressure (bar)
90
30
6
2.3
Enthalpy (MJ/kmol)
58.6
50.5
49.7
49
Table 7-1 The heat content of the steam at different levels.
199
of the major parameters in the consideration of the economics
of the utility system.
The two boilers use the natural gas as a fuel. This gas is
one of the most commonly used fuels, because it is cheap
compared with other fuels, clean doesn't leave ash after
burning and is easy to handle. The net calorific value for
natural gas is 37.670 MJ/m3 .
Both boilers receive saturated water at 100 °C to be
heated and pressurized to 450 °C and 90 bar respectively. The
enthalpy of saturated water at 100 °C is equal to 7.54
MJ/kmol.
The heat taken by the working fluid in each boiler = heat
content in the steam leaving the boiler - heat content of the
saturated water entering the boiler = 7.15 (58.6 - 7.54) =
365.079 MW.
The boiler efficiency is defined by the following
equation;
Heat transferred to working fluid Boiler efficiency = ——
Boiler efficiency =
Fuel energy supplied
Heat transferred to working fluid
Fuel consumption x net calorific value
The boiler efficiency is taken to be 85% (allowance for
200
loss of heat through walls and stack), therefore the amount
of fuel consumed in each boiler will be;
365.079 0.85 = ————————————
Fuel consumption x 37.67
Fuel consumption - 11.4 m3 /s
therefore the total amount of fuel consumed in the boilers is
22.8 m3 /s.
7.1.2 Application of a Combined Heat and Power System in the
Styrene Plant After Energy Integration
Energy integration within the styrene process has revealed
a significant saving in energy and led to a great reduction
in utility imports. The most important result derived from
this energy integration within the styrene process design is
that, the heat demand of all reboilers associated with
distillation trains in both the alkylation and
dehydrogenation processes are satisfied by interprocess heat
exchange. This means that the styrene process no longer needs
to import steam at low and high pressures for process
heating. A reduction of 4.7 kmol/s in steam import for the
process has been achieved. The result of this reduction has
decreased the amount of fuel consumption by 7.5 m3 /s, and
this amounts to the third of the total fuel originally used
201
in the boiler house.
The rest of the steam import (which accounts to 9.6 m3 /s)
is required by the dehydrogenator and is at intermediate
pressure. The styrene process after being energy integrated
is capable of producing 7.5 kmol/s of steam at intermediate
pressure. Thus this steam will replace most of the
intermediate pressure steam imported from the boiler house.
Therefore the process at this stage needs only 2.1 kmol/s of
intermediate pressure steam to be imported. This has resulted
in a reduction of 11.95 m3 /s in the amount of fuel consumed,
leaving the process with a need of only 3.35 m3 /s. Not only
is this significant amount of fuel saved, but the great
reduction in requirements of steam has led into a complete
shutdown of one of the boilers, also the size of the other
boiler can be reduced to less than half. Therefore the
boilers capital cost, maintenance, operators, and other fixed
costs are also saved.
In contrast, the amount of power generated in the plant
will be decreased due to the small amount of steam passing
through the central turbine. Consequently much a smaller
turbine is needed.
The amount of work available due to the steam that is
subsequently used by the dehydrogenator is;
W = 2.1 x 58.6 - 2.1 x 49.68 = 18.7 MW
202
thus the total electricity produced in the plant will be
about 15 MW. The estimate power demand for the process
including the compressers added in the design process is of
the order of 25 MW< 84 >. Therefore the rest of the plant needs
for electricity will be taken directly from the electricity
board. The connection of the heat and power system with the
process after integration is illustrated in Figure 7-2.
7.2 Concluding Remarks
The energy integration of the styrene process has resulted
in a great reduction in utility consumption. Also steam is
raised in the process by recovering all surplus heat in a
waste heat boiler. This steam is used to cover most of the
process steam demands, resulting in a very small need for the
utility to be imported from the boiler house. Therefore, one
of the boilers is completely shutdown, and the fuel
consumption is dramatically reduced.
203
ruel—— •"! —
11.4 ™I
Y/ 7.15vkmo111 rvH n. r » YDOWei /\
/ 1\ "
S / \ V
Fuel^^
3-35 ™?s
ykmolBoiler ^~
2
Standby L_Boiler
______ >/
EoV
I/IQ.
_LO
i
0Eui
JC
q
C O l/) ^
2 1 k mo1 steam at 450° s and 90 bar
^^Power station-^L
f\ o
ID•oC0
0oto
oE0
I/)a.
1
° s
*-^
\/
oE«/l
^e.
N
1
*— -
———— 0 w.
••-».
/SiV W S\w*i
Q.
^O
"o
oEt-.r>4
CgC ««a» V
?§^ P "° o* ^ — s ——————— £< f Q;
/ . » Q
//
15MW
Figure 7.2
The heat exchange between dehydrogenation and alkylation
The boiler house and the power station 4n connection with styrene process after energy integration.
204
CHAPTER EIGHT
Conclusions
The conclusions drawn from this research work can be
summarized as follows;
1- Process energy integration does not necessarily mean the
use of advanced equipment, since the utility import
depends greatly on the way that process streams are linked
(via heat exchangers).
2- The number of distillation sequences to be evaluated can
be reduced using the stated heuristics. Therefore the
optimum or near optimum distillation sequence can be found
without the need for testing all possible sequences. The
selected sequence can then be energy integrated with the
rest of the process streams.
3- In order to maintain the best links between process
streams, each stream has to be taken in the context of the
overall process, so that all possibilities for heat
exchange with the specified stream are considered.
4- The design that reaches the process energy targets is not
necessarily the best design. Since opportunities for
better network design can always arise if improvements are
sought in terms of process stream links or design
205
parameters.
5- The reduction of fuel consumed in the process is not
always cost effective, especially when the process
produces a by product that can be used as a fuel within
the process. It has been shown in this research work that
increasing the use of such by product fuel can reduce the
import of fuel dramatically else vAere in the plant. Also
the process network design can be greatly improved
resulting in a significant decrease in the import of
utility.
6- The process can always benefit from cold utilities at
higher temperature levels, since further utility
generation can then be sought. This has two positive
effects on the process which are decreasing the demand of
the process for utility, and ensuring the availability of
the utility at no fuel cost.
7- The utility system should not be designed in isolation
from the process.
8.1 Future Investigations
As the energy consumption in a process design decreases
the capital cost is likely to increase. Therefore one of the
206
aims for the future investigations is to find the balance
between the decrease in energy cost and the increase in the
capital cost. Thus finding the global optimum process in
terms of energy and capital.
Designs based upon smaller driving forces in the process
will always reduce the operational flexibility of the system
and consequently make the process difficult to control.
Therefore, process flexibility and control are suggested as
one of the areas for future investigation.
As the ultimate design has to be reliable, then this
subject is also recommended for investigation in the future.
207
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212
34- Townsend D,W. and Linnhoff B., "Surface Area Targets for
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35- Ahmed S. and Linnhoff B., "Overall Cost Targets for Heat
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36- Lecture 7, Short Course on Pinch Technology, University
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37- Linnhoff B. and Vredeveld D.R., "Pinch Technology Has
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213
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44- Doukas N. and Luyben W.L., "Economics of Alternative
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45- Seader J.D. and Westerberg A.W., "A Combined Heuristic
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47- Thomas R.W. and King C.J., "Systematic Synthesis of
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48- Hendry J.E. and Hughes R.R., "Generating Separation
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49- Nath R. and Motard R.L., "Evolutionary Synthesis of
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50- Westerberg A.W. and Andrecovich M.J., "Utility Bounds for
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51- Petterson W.C. and Thomas A.W., "Energy Saving Schemes in
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52- Rathore R.N.S., Vanwormer K.A. and Powers G.J.,
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53- Tedder D.W, and Rudd D.F., "Parametric Studies in
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54- Stupin W.J. and Lockhart F.J., "Thermally Coupled
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55- Andrecovich M.J. and Westerberg A.W., "A Simple Synthesis
Method Based on Utility Bounding for Heat Integrated
215
Distillation Sequences", AIChE J, vol.31, p.363 (1985).
56- Flower J.R. and Jackson R. , "Energy Requirements in the
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62- Linnhoff B. and Townsend B.W., "Designing Total Energy
216
Systems", Chem. Eng. Prog,, vol.78, p.72 (1982).
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64- Townsend D.W. and Linnhoff B., "Heat and Power Networks
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Engines and Heat Pumps in Process Networks", AIChE J,
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65- Townsend D.W. and Linnhoff B., "Heat and Power Networks
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217
of Heat Integrated Distillation Sequences", IChemE Symp.
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70- Muraki M. and Hayakawa T., "Evolutionary Synthesis Method
of Energy Integrated Distillation Separation Process",
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72- King C.J., "Separation Processes", 2nd Edition, MacCrow
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73- Fonyo Z., Meszaros I., Rev E. and Kaszas M. , "Pinch
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Engineering" Volume 6, 1st Edition (1983).
75- Lecture 5, Short Course on Pinch Technology, The
University of Manchester Institute of Science and
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76- Short Course on Pinch Technology, Part 2, The University
218
of Manchester Institute of Science and Technology, Jan.
9-12 (1989).
77- Huang W. , "Operation Simulation Analysis Boosts Styrene
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78- CdF Chimie-Technip, Hydrocarbon Processing, p.244 (Nov.
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79- Linnhoff B. , "The Process/Utility Interface", 2nd
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(Mar. 1986).
80- Balzhiser R.E., Samuels M.R. and Eliassen J.D., "Chemical
Engineering Thermodynamics", (1972).
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Properties of Gases and Liquids" 3rd Edition, McCraw Hill
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82- Glassman I., "Combustion", (1977).
83- Spires H.M., "Technical data on Fuels", 6th Edition,
Published by the British National Committe World Power
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219
84- Personnal discussions with industrial contacts, C.D.G.,
1990.
220
- APPENDIX A -
MASS AND ENERGY BALANCE SAMPLE CALCULATION
A.I Mass Balance Sample Calculation
This mass balance sample calculation is for the first
distillation column in the dehydrogenation process. This
column is called the benzene-toluene column, and it separates
the benzene and toluene from the crude styrene as an overhead
product.
Overall mass balance:
F = D + B
F = crude styrene = 1779.7 kmol
D = benzene + toluene
= 14.374 + 22.34 = 36.7 kmol
B = F - D
B = 1779.7 - 36.7 = 1743 kmol
Reflux ratio (R) = 12
As R = L / D
Then L = R x D
L = 12 x 36.7 = 440.4 kmol
221
A, 2 Energy Balance Sample Calculation
The energy balance sample calculation is for the
alkylator. The reaction in this reactor is represented by
the following equation;
95°C2H4 + CeHe ———————*• CsHio Et Bz EB
The base temperature is taken to be 25 °C (the standard
state for heat of reaction). The thermodynamic pathway for
determining the total heat absorbed (or released) in the
alkylator is shown in Figure A-l. The reactants are cooled
from 40° to 25 °C, the reaction is carried out at 25 °C and
the products are then heated from 25 °C to the reaction
temperature (95 °C).
Q = S AHp + QR - S aHR ———————————————————— (A-l)
where S AHn is the enthalpy change to bring the reactants to
the standard temperature, S <aHp is the enthalpy change to
bring the products to the reaction temperature and QR is the
total heat generated by the reaction taking place, evaluated
from the standard heat of reaction at 25 °C.
All enthalpy values are found by using the physical
properties data service (ppds) program except the
triethylbenzene , since this component is not included in the
program. Therefore the group-contribution method of Luria
222
Et = 907.9 kmol Bz = 1510.9 kmol
DEB = 142.3 kmolTEB = 7 kmol
40°______
25°
95°
Bz = 891.08 kmol EB = 721.83 kmol
DEB = 145.17 kmolTEB =7.4 kmol
Figure A-l A thermodynamic pathway to determine the netheat absorbed or released from the alkylator.
and Benson* 8 1 > is used to calculate the heat capacity of the
triethylbenzene as follows;
Type of group
CB - (H)
CB - (C)
C - (CB )(C)(H)2
C - (C)(H) 3
No. of groups
3
3
3
3
CP = 3 (-1.842 + 0.05788 T - 0.0001716 T2 + 1.995x10' 7 T3 )
+ 3 (28.807 - 0.2824 T + 0.0009779 T2 - l.lOSxlO' 6 T3 )
+ 3 (30.192 - 0.2812 T + 0.001002 T2 - 1.1150xlO- 6 T3 )
+ 3 (8.459 + 0.002113 T - 5.605xlO- 5 T2 + 1.723x10- 7 T3 )
= 196.848 - 1.5111 T + 0.000525675 T2 - 5.5386xlQ- 6 T3
223
= - 598.124 MJ
/»Hu 2 = - 3109.432 MJ
£>HDEB = - 516.122 MJ
AHTEB = - 31.657 MJ
Therefore 2 AHn = - 4255.335 MJ
2 AHp = &HBZ + &HEB + &HDEB + &HTEB
= 9016 MJ
= 10092. 627MJ
= 2585.46 MJ
= 157.54 MJ
Therefore S 6Hp = 21851.627 MJ
The standard heat of reaction is equal to the heat of
formation of the products minus the heat of formation of the
reactants .
&H.2S = 2 £HF(prod.) - 2 6HF(reac.)
Heat of formation of ethylbenzene (EB) = - 12.503 MJ/kmol
Heat of formation of ethylene (Et) = 52.483 MJ/kmol Heat
of formation of benzene ( Bz ) = 48.846 MJ/kaol
Therefore &H25 = - 113.832 MJ/kmol
The minus sign indicates that the reaction is exothermic.
QR = AHzs x mols produced
QR = - 113.832 x 721.83 = - 82167.35 MJ
224
Thus the heat released by the reactor will be;
Q = 21851.627 - 82167.35 - 4255.335 = - 64571.058 MJ
225
- APPENDIX B -
VAPOUR PRESSURE, RELATIVE VOLATILITY AND AVERAGE RELATIVE
VOLATILITY DATA
B.I Vapour Pressure Values
Column No.
1
2
3
Component Name
Benzene
EB
DEB
TEB
EB
DEB
TEB
DEB
TEB
V.P. at the Top N/m2
104194
17422.2
3235.2
1128.2
106344
28262
11222.7
104374
45607.3
V.P. at Feed Stream N/m2
24379.6
2873
364.5
118
137694
38432.6
15598.4
132443
58909
V.P. at the Bottom N/m2
558646
137694
38432.6
15598.4
391732
132443
58909
228791
105920
Table B-l Vapour pressure values for each component ineach column in the sequence shown in Figure 4-la,
226
Column No.
Component Name
V.P. at the Top N/m2
V.P. at Feed Stream N/m2
V.P. at the Bottom N/m2
The values of the first column are the same as in Table B-l
2
3
EB
DEB
TEB
EB
DEB
TEB
148379
42002
17156.8
106344
28262
11222.7
137694
38432.6
15598.4
148379
42002
17156.8
623979
228791
105920
391732
132443
58909
Table B-2 Vapour pressure values for each component ineach column in the sequence shown in Figure 4-lb,
227
ColumnNo.
1
2
-
Component Name
Benzene
EB
DEB
TEB
Benzene
EB
- -V.P. at the
Top N/m2
278789
58737.6
13917.3
5269.2
104194
17422.2
V.P. at Feed Stream N/m2
24379.6
2873
364.5
118
278789
58737.6
V.P. at the Bottom N/m2
1325000
391732
132443
58909
524935
127618
The values of third column are the same as in Table B-l
Table B-3 Vapour pressure values for each component ineach column in the sequence shown in Figure 4-lc.
228
Column No.
1
2
3
Component Name
Benzene
EB
DEB
TEB
Benzene
EB
DEB
TEB
EB
DEB
TEB
V.P. at the Top N/m2
396111
90420.5
23295.2
9127
104194
17422.2
3235.2
1128.2
106344
28262.4
11222.7
V.P. at Feed Stream N/m2
24379.6
2873
364.5
118
396111
90420.5
23295.2
9127
137694
38432.6
15598.4
V.P. at the Bottom N/m2
1957000
623979
228791
105920
558646
137694
38432.6
15598.4
391732
132443
58909
Table B-4 Vapour pressure values for each component ineach column in the sequence shown in Figure 4-ld.
229
Column No.
Component Name
V.P. at the Top N/m2
V.P. at Feed Stream N/m2
V.P. at the Bottom N/m2
The values of first column are same as the values of column No.l in Table B-4
Benzene
EB
DEB
TEB
278789
58737.6
13917.3
5269.2
396111
90420.5
23295.2
9127
1325000
391732
132443
58909
The values of third column are same as the values of column No.2 in Table B-3
Table B-5 Vapour pressure values for each component ineach column in the sequence shown in Figure 4-le.
230
B.2 Relative Volatility Values
Column No.
1
2
3
Component Name
Benzene
EB
DEB
TEB
EB
DEB
TEB
DEB
TEB
R.V. at the Top
92.35
15.44
2.8
1
9.47
2.5
1
2.28
1
R.V. at Feed
206.6
24.34
3.1
1
8.8
2.46
1
2.24
1
R.V. at the Bottom
35.8
8.8
2.46
1
6.6
2.25
1
2.16
1
Table B-6 The relative volatility values of each component in each column in the sequence shown in Figure 4-la.
231
ColumnNo.
Component Name
R.V. at the Top
R.V. at Feed R.V. at the Bottom
The values of first column are the same as the values of first column in Table B-6
2
3
EB
DEB
TEB
EB
DEB
TEB
8.65
2.45
1
9.5
2.5
1
8.8
2.46
1
8.6
2.45
1
6
2.16
1
6.6
2.25
1
Table B-7 The relative volatility of each component ineach column in the sequence shown in Figure 4-lb,
232
Column No.
1
2
3
Component Name
Benzene
EB
DEB
TEB
Benzene
EB
DEB
TEB
R.V. at the Top
53
11
2.6
1
6
1
2.28
1
R.V. at Feed
206.6
24.34
3.1
1
4.7
1
2.24
1
R.V. at the Bottom
22.5
6.6
2.25
1
4
1
2.16
1
Table B-8 The relative volatility of each component ineach column in the sequence shown in Figure 4-lc.
233
Column No.
1
2
3
Component Name
Benzene
EB
DEB
TEB
Benzene
EB
DEB
TEB
EB
DEB
TEB
R.V. at the Top
43.4
10
2.5
1
92.35
15.44
2.86
1
9.47
2.5
1
R.V. at Feed
206.6
24.34
3.1
1
43.4
10
2.5
1
8.8
2.46
1
R.V. at the Bottom
18.5
5.8
2.16
1
35.8
8.8
2.46
1
6.65
2.25
1
Table B-9 The relative volatility of each component ineach column in the sequence shown in Figure 4-ld.
234
Column No.
Component Name
R.V. at the Top
R.V. at Feed R.V. at the Bottom
The values of first column are same as the values of first column in Table B-9
2
Benzene
EB
DEB
TEB
53
11
2.6
1
43.4
10
2.5
1
22.5
6.6
2.25
1
The values of third column are same as the values of second column in Table B-8
Table B-10 The relative volatility of each component in each column in the sequence shown in Figure 4-le.
235
B.3 Average Relative Volatility
Sequence (a)
Sequence(b)
Sequence (c)
Sequence (d)
Sequence (e)
Component Name
Benzene
EB
DEB
TEB
Benzene
EB
DEB
TEB
Benzene
EB
DEB
TEB
Benzene
EB
DEB
TEB
Benzene
EB
DEB
TEB
Column No.l
88
14.9
2.8
1
88
14.9
2.8
1
62.6
12
2.6
1
55
11.2
2.6
1
55
11.2
2.6
1
column No. 2
8.2
2.4
1
7.65
2.35
1
4.86
1
52.3
11
2.6
1
37.2
9
2.4
1
Column No. 2
2.2
1
8.15
2.39
1
2.2
1
8.2
2.4
1
4.8
1
Table B-ll The average relative volatility for each component in each column in the sequences shown in Figure 4-1.
236
- APPENDIX C -
COMPUTER PROGRAM FOR CALCULATING THE ROOT OF UNDERWOODS
EQUATIONS
cC * Program For Determining the Underwood Parameter {Theta) *
Real*8 Th(lOOOO), Det(lOOOO) Dimension Alpha(lOO), Xf(lOO)
C
800 Print*,' Number of Components, N ?, and q ? '
Read*, N, q
C
C N = Number of components
C q = State of feed
C
Do 1 k = 1, N
Print*,' Value of Alpha(I), Xf(I), I = f ,k
Read*, Alpha(k), Xf(k)
C
C Alpha(k) = Relative volatility of the kth component
C Xf(k) = Mole fraction of the kth component in the feed
C
1 Continue
C
1000 Print*,' Initial Value of Theta?, Final Value of Theta'
Read*, Thetai, Thetaf
C
Xinc = (Thetaf - Thetai )/100
9 Ic = 0
Do 10 Theta = Thetai, Thetaf, Xinc
C
C Theta = Underwoods Parameter
237
Ic = Ic + 1
Sum2 = 0
Sum3 = 1
Do 20 j = 1, N
Suml = 1
Do 30 i = 1, N
If (J.eq.l) Sum3 = Sum3 * (Alpha(I) - Theta)
If (i.eq.j) GOTO 30
Suml = Suml * (Alpha(I) - Theta)
30 Continue
Sum2 = Sum2 + Suml * (Alpha(j) * Xf(j))
20 Continue
Sum4 = Sum3 * (q - 1)
Det(Ic) = Sum2 + Sum4
Th(Ic) = Theta
C
If (Abs (Det(Ic)).LE.0.01) Then
kO = 1
GOTO 200
ENDIF
Print*,' Det, Th(Ic), 1=', Det(Ic), Th(Ic), Ic
C
If (Ic.GE.3) Then
Detl = Abs(Det(Ic-2) )
Det2 = Abs(Det(Ic-l) )
Det3 = Abs(Det(Ic))
If (Det2.LE.Detl.and.Det2.LE.Det3) Then
238
Print*,' Test '
Thetaf = Th(Ic)
Thetai = Th(Ic-2)
Thl = Abs(Th(Ic))
Th2 = Abs(Th(Ic-l))
If((Thl - Th2).LE.0.01) GOTO 300
Xinc = (Thetaf - Thetai)/100
If(Xinc.LE.O) GOTO 300
GOTO 9
CENDIF ENDIF
C
10 Continue
C
200 If (KO.EQ.l) Then
Print*,' Theta =', Th(Ic),' Det=', Det(Ic)/Sum3
KO = 0
GOTO 400
ENDIF
C
C
300 Detl = Det(lc-l)/Sum3
Print*,' Theta=', Th(Ic-1),'Det=',Detl
C
400 Print*,'Do You Like to Re-start the Programme, Type 1'
Bead*, yes
If (yes.EQ.l) GOTO 800
239
Print*,'Do You Like to Have Anather Range For Theta'
Print*,' If Yes Type 1'
Read*, yes
If(Yes.EQ.l) GOTO 1000
End
240
- APPENDIX D -
THE DATA FOR CALCULATING THE OPTIMUM REFLUX RATIO
D.I The Sequence Shown in Figure 4-la
Column No.l:
Minimum number of plates (Sm) = 6.75
Minimum reflux ratio (Rm) = 0.3
R» / R« + 1 = 0.23
Times Rn
1
1.05
1.08
1.1
1.2
1.4
1.7
2
2.5
2.75
3
6
9
15
20
R
0.3
0.315
0.324
0.33
0.36
0.42
0.51
0.6
0.75
0.825
0.9
1.8
2.7
4.5
6
R/R+1
0.23
0.239
0.244
0.248
0.265
0.295
0.337
0.375
0.428
0.452
0.47
0.642
0.729
0.818
0.857
s./s0
0.04
0.11
0.12
0.27
0.39
0.52
0.6
0.66
0.69
0.715
0.84
0.9
0.935
0.955
S
CD
168.7
61.3
56.25
25
17.3
13
11.25
10.22
9.78
9.44
8
7.5
7.2
7
Table D-l The data for estimating the optimum reflux ratioof column No.l in the sequence shown in Figure 4-la,
241
Column No.2 :
Minimum number of plates (Sn) = 8.2
Minimum reflux ratio (RB ) = 0.44
Rm/R«+l = 0.3
Times Ra
1
1.05
1.08
1.1
1.3
1.4
1.7
2
2.5
2.75
3
6
9
15
20
R
0,44
0.462
0.475
0.484
0.57
0.616
0.748
0.88
1.1
1.2
1.32
2.64
3.96
6.6
8.8
R/R+1
0.3
0.316
0.322
0.326
0.363
0.38
0.428
0.468
0.524
0.547
0.569
0.725
0. 798
0.868
0.897
S«/S
0
0.11
0.2
0.24
0.41
0.49
0.62
0.68
0.75
0.78
0.79
0.895
0.928
0.96
0.97
S
<a
74.5
41
34.2
20
16.73
13.2
12
10.93
10.5
10.38
9.16
8.84
8.54
8.45
Table D-2 The data for estimating the optimum reflux ratioof column No.2 in the sequence shown in Figure 4-la,
242
D.2 The Sequence Shown in Figure 4-lb
The conditions of column No. 1 in this sequence are the
same as the conditions of column No.l in the sequence shown
in Figure 4-la, thus the data for estimating the optimum
reflux ratio are same as the ones in Table D-l.
Column No.3 :
Minimum number of plates (Sm) = 8.2
Minimum reflux ratio (R« ) = 0.66
Rn / R» + 1 = 0.4
Times Rm
1
1.05
1.08
1.1
1.3
1.4
1.7
2
2.5
2.75
3
6
R
0.66
0.693
0.713
0.726
0.86
0.924
1.12
1.32
1.65
1.815
1.98
3.96
R/R+1
0.4
0.41
0.416
0.42
0.46
0.48
0.528
0.57
0.62
0.644
0.664
0.76
Sm/S
0
0.1
0.22
0.28
0.465
0.525
0.66
0.71
0.79
0.818
0.822
0.9
S
00
82
37.27
29.28
17.63
15.6
12.4
11.55
10.38
10
9.97
9.11
Table D-3 The data for estimating the optimum reflux ratioof column No.3 in the sequence shown in Figure 4-lb.
243
D.3 The Sequence Shown in Figure 4-lc
Column No.l :
Minimum number of plates (Sm) = 6.55
Minimum reflux ratio (Rm) = 0.13
R./R.+ IB 0.115
Times Rn>
1
1.05
1.08
1.1
1.4
1.7
2
2.5
3
6
10
15
20
30
40
R
0.13
0.136
0.14
0.143
0.182
0.221
0.26
0.325
0.39
0.78
1.3
1.95
2.6
3.9
5.2
R/R+1
0.115
0.12
0.123
0.125
0.154
0.18
0.2
0.245
0.28
0.438
0.565
0.66
0.72
0.79
0.83
Sm/S
0
0.04
0.045
0.05
0.2
0.315
0.4
0.46
0.5
0.67
0.79
0.86
0.89
0.93
0.95
S
00
163.75
145.55
131
32.75
20.8
16.35
14.24
13.1
9.77
8.3
7.6
7.36
7
6.89
Table D-4 The data for estimating the optimum reflux ratio of column No.l in the sequence shown in Figure 4-lc
244
Column No.2 :
Minimum number of plates (Sm) = 7.7
Minimum reflux ratio (Km) = 0.5
Rni / Rm + 1 =0.33
Times Rn
1
1.05
1.08
1.1
1.2
1.4
1.7
2
4
6
8
10
20
R
0.5
0.525
0.54
0,55
0.6
0.7
0.85
1
2
3
4
5
10
R/R+1
0.33
0.344
0.35
0.354
0.375
0.4
0.46
0.5
0.66
0.75
0.8
0.83
0.9
Sm/S
0
0.1
0.15
0.22
0.37
0.475
0.6
0.72
0.86
0.92
0.935
0.94
0.967
S
CD
77
51.3
35
20.8
16.2
12.42
10.69
8.95
8.369
8.235
8.2
7.9
Table D-5 The data for estimating the optimum reflux ratioof column No.2 in the sequence shown in Figure 4-lc.
245
D.4 The Sequence Shown in Figure 4-ld
Column No.2 :
Minimum number of plates {So) = 7.27
Minimum reflux ratio (Rm) = 1.44
Rm / Em + 1 = 0.59
Times Rm
1
1.05
1.08
1.1
1.2
1.4
1.7
2
3.5
5
7.5
10
20
R
1.44
1.512
1.555
1.584
1.7
2.106
2.448
2.88
5.04
7.2
10.8
14.4
28.8
R/R+1
0.59
0.6
0.608
0.613
0.63
0.668
0.709
0.742
0.834
0.878
0.915
0.935
0.966
SB /S
0
0.18
0.3
0,32
0.4
0.52
0.66
0.725
0.84
0.89
0.92
0.94
0.96
S
00
40.38
24.2
22.7
18.17
13.98
11
10
8.65
8.16
7.9
7.7
7.57
Table D-6 The data for estimating the optimum reflux ratioof column No.2 in the sequence shown in Figure 4-ld.
Column No. 3 : The conditions of this column are the same
as the conditions of column 2 sequence a.
246
D.5 The Sequence Shown in Figure 4-le
Column No.2 :
Minimum number of plates (Sn) = 7.7
Minimum reflux ratio (Rm) = 0.35
Rn / R« + 1 = 0.26
Times R»
1
1.05
1.08
1.1
1.3
1.4
1.7
2
2.5
3
5
7.5
10
20
R
0.35
0.367
0.378
0.385
0.445
0.49
0.595
0.75
0.875
1.05
1.75
2.625
3.5
7
R/R+1
0.26
0.268
0.274
0.278
0.3
0.328
0.37
0.4
0.466
0.5
0.636
0.724
0.777
0.875
S./S
0
0.07
0.1
0.16
0.285
0.35
0.52
0.6
0.7
0.74
0.84
0.89
0.925
0.96
S
00
110
77
48.12
27
22
14.8
13
11
10.4
9.16
8.65
8.3
8
Table D-7 The data for estimating the optimum reflux ratio of column No.2 in the sequence shown in Figure 4-le
Column No.3 : The conditions of this column are the same as ones of column 3 sequence c.
247
D.6 The Sequence Shown in Figure 4-10c
Column No.l :
Minimum number of plates (Sm) = 14.9
Minimum reflux ratio (Km) = 7.86
Rm / Rm + 1 = 0.887
Times Rm
1
1.05
1.08
1.1
1.3
1.5
1.6
2
2.5
3
10
15
20
R
7.86
8.25
8.487
8.64
10.2
11.788
12.57
15.718
19.64
23.577
78.59
117.89
157.18
R/R+1
0.887
0.89
0.894
0.896
0.91
0.92
0.926
0.94
0.95
0.96
0.98
0.99
0.994
SB /S
0
0.03
0.09
0.11
0.4
0.525
0.64
0.73
0.77
0.8
0.94
0.97
0.99
S
<D
496.6
165.5
135.45
37.25
28.38
23.28
20.4
19.35
18.625
15.85
15.2
15
Table D-8 The data for estimating the optimum reflux ratioof column No.l in the sequence shown in Figure 4-10c,
248
Column No.2 :
Minimum number of plates (Sm) = 35.8
Minimum reflux ratio (Rm) = 5
Rn / R« + 1 = 0.83
Times Rm
1
1.05
1.08
1.1
1.23
1.4
1.7
2
2.5
3
10
15
20
R
5
5.25
5.4
5.5
6.15
7
8.5
10
12.5
15
50
75
100
R/R+1
0.83
0.84
0.843
0.846
0.86
0.875
0.894
0.91
0.926
0.937
0.98
0.986
0.99
Sn,/S
0
0.14
0.16
0.18
0.45
0.5
0.68
0.75
0.8
0.825
0.95
0.96
0.98
S
00
255.7
223.75
198.8
79.5
71.6
52.64
47.7
44.75
43.4
37.68
37.3
36.5
Table D-9 The data for estimating the optimum reflux ratio of Column No.2 in the sequence shown in Figure 4-10c
249
Column No.3 :
Minimum number of plates (Sa) = 12
Minimum reflux ratio (Rm) = 1.6
Rm / Rm + 1 = 0.616
Times Rm
1
1.05
1.08
1.1
1.25
1.4
1.7
2
2.5
3
5
10
15
20
R
1.6
1.684
1.73
1.764
2
2.245
2.727
3.208
4.0
4.812
8.02
16.04
24.06
32.08
R/R+1
0.616
0.627
0.63
0.638
0.66
0.69
0.73
0.76
0.8
0.828
0.89
0.94
0.96
0.97
Sm/S
0
0.15
0.18
0.22
0.52
0.55
0.66
0.71
0.78
0.815
0.88
0.94
0.95
0.98
S
00
80
66.6
54.54
23
21.8
18.18
16.9
15.4
14.7
13.6
12.76
12.6
12.24
Table D-10 The data for estimating the optimum reflux ratio of column No.3 in the sequence shown in Figure 4-10c
250
2 ^ in <b o Q. O. o ^ 0> £ 3 z;
500
475;
450^
425:
400
375.
350:
325.
300:
275.
250
:
225.
200.
175
150
125
100 75 50 25
0L. .
0.0
01.
70
3.4
05.
10
6.80
8.
50
10.2
0 R
eF
Lux
ratio
<
R)
1.90
13
.60
15.3
0 X1
01
17.00
Figure D.I
The
rela
tion
be
twee
n re
flux
ra
tio
and
numb
er of
plates in
the
firs
t column of
sequence (c)
in
Figu
re 4-10.
ro 01
to
Cb d 0.
0. o i. -D £ 3
300
285i
270.
255-
240
225
210.
195.
180
165
150
135
120
105 90 75 60 45 30 15 0
T-I
-p
10
20
30
40
50
60R
eFLu
x ra
tio (R
)70
8090
100
Figu
re D.
2 The
rela
tion
be
twee
n reflux ratio
and
numb
er of
plates in
the
second column of
se
quen
ce (c
) in
Figure 4-10.
(Jt
in d -p o a a_ o i. Q> £1 e 15
100 95 \
90:
85:
so:
75.
70.
65.
60.
55.
50.
45:
40:
35:
30-
25:
20:
15 10. 5: 0. 0.0
0
0.3
5
0.7
0I.
05
I.40
1.75
2.
10
Re
FLux
ratio
(R
)2.45
2.80
3.15
3.50
X10
1
Figure D.
3 The
relation be
twee
n reflux ratio
and
number of
plates in the
thir
d co
lumn
of sequence (c)
in
Figure 4-10.
- APPENDIX E -
FLAME TEMPERATURE CALCULATION
E.I The Flame Temperature Calculation
The flame temperature differs from one fuel to the other,
and different parameters can effect this temperature such as,
the compositions of the fuel, the amount of excess air that
is used and whether this air is preheated or not.
As mentioned above the fuel used in the furnace is the gas
that is vented from the process as a by-product. This gas
consists mainly hydrogen (over 85%), and some other gases
such as, methane and very small percentage of carbon
monoxide, carbon dioxide and ethane. For simplicity the
compositions taken to be 85% of hydrogen and 15% of methane
(both percentages are by volume). The heats of oxidation are
242 kJ/mol and 802 kJ/mol for hydrogen and methane
respectively(82 ) . The stoichiometric amount of oxygen
required to satisfy the combustion reactions is calculated as
follows;
Ha + 1/2 Oz ———————*> HaO AH = - 242 kJ/mol 0.85 0.425 0.85
CH4 + 2 02 ———————»• CO2 + 2 H20 £>H = 802.8 kJ/mol 0.15 0.3 0.15 0.3
The combustion of 1 mole of the gas is taken as a basis.
Therefore the oxygen required = 0.725 x 0.0224 = 0.0163 m3
and thus the amount of air required = 0.0163 x 100/21 = 0.077
m3 , the associated amount of nitrogen = 0.077 - 0.0163 = 0.06
254
m3 .
The product of the combustion is assumed to carry an
excess of 8 per cent by volume of unburned oxygen. Therefore
the product will consist of 0.0034 m3 COz , 0.0258 m3 HaO,
0.087 m3 N2 and 0.007 m3 02 = 0.123 m3 . The specific heats of
COz, HzO, N2 and 02 are respectively 2.45, 1.95, 1.512 and
1.58 kJ/m3 /°C< 83 > . These are mean values between 0° and 2000
Therefore the mean specific heat of the combustion gases =
(0.0034/0.123) 2.45 + (0.0258/0.123) 1.95 + (0.087/0.123)
1.512 + (0.007/0.123) 1.58 = 1.64 kJ/m3 / 0 C
326.12The temperature rise = —————————————— = 1616.7 ° C
0.123 x 1.64
After adding the ambient temperature of 20 °C, the flame
temperature will be 1636.7 °C.
Assuming 5 per cent heat loss during combustion the
temperature rise will be 1536 ° C . Therefore the flame
temperature would be 1556 °C. For this study the flame
temperature through calculations is taken to be 1500 *C.
255
