CPD NR 3258
Conceptual Process Design
Process Systems Engineering
DelftChemTech - Faculty of Applied Sciences
Delft University of Technology
Subject
Dehydration of ethanol
Authors
M. de Jong
P.F.A. van Rooijen
V. Verboom
T. Winkels
Telephone
0180-419065
015-3108063
015-2132840
0174-295760
Keywords
Ethanol, dehydration, water, separation, distillation,
membranes
Assignment issued
Report issued
Appraisal
:
:
:
09-03-2001
12-06-2001
26-06-2001
Summary
Ethanol is one of the main base chemicals in the world and is often used as a fuel
additive to produce gasohol. Before the ethanol can be used as additive it needs to be
dehydrated and purified up to 99.8 vol%. This report provides a comparison between
four conceptual designs of ethanol dehydration units, which process a feedstream of
88.8 vol% ethanol to the required purity of at least 99.8 vol%. During the process
impurities like ethyl acetate, acetaldehyde, isobutyl alcohol and isopentyl alcohol are
partly removed. The alternatives, azeotropic distillation by toluene, extractive
distillation by gasoline, extractive distillation by polyacrylic acid (PAA) and normal
distillation followed by membrane purification, are chosen out of a wide range of
options found in literature. After comparison of the four alternatives on the basis of
validity of thermodynamic data, safety, environmental impact and economy, the
normal distillation followed by membrane purification appears to be the most
interesting option for further design.
The ethanol dehydration plant is intended to use a specific feed stream originating
from a nearby ethanol fermentation plant. Therefore the production capacity of the
dehydration plant is fixed to 8.6 kton (10.9 kliters) of ethanol per year. This is in
comparison with the production level in Europe of 4.7 billion liters per year, a very
small amount. Therefore the sales price for ethanol produced in the dehydration plant
will be dependent on the world market price. The dehydration plant is a continuous
operated plant, which is 8,620 hours per year on stream (stream factor 0.984). This
implies that 150 hours per year are accounted for unexpected shut-downs. The plant
life is 15 years, which includes 2 years of construction and 1 year of deconstruction.
The options azeotropic distillation by toluene, extractive distillation by gasoline and
normal distillation followed by membrane purification are already operated at full
scale and are patented. The extractive distillation by PAA is based on laboratory
experiments. During the design some doubts arose for the thermodynamic properties
of this last option and therefore the extractive distillation by PAA is rejected in the
selection procedure for the final design.
The total investment costs are the highest for the azeotropic distillation by toluene
with 1.8 million EURO. The costs of extractive distillation by PAA and normal
distillation followed by membrane purification are modest with respectively 1.2
million EURO and 0.9 million EURO. The option extractive distillation by gasoline
has the lowest investment of 0.7 million EURO. Calculation of the economic criteria
shows that the azeotropic distillation by toluene is not profitable. This option has a
negative cash flow of 0.9 million EURO per year and is therefore rejected as a final
design option. In reviewing the variable costs it appeared that at least 58 % of the
operating costs are caused by the purchase of the raw materials. Therefore the
sensitivity of the designs for changes of 5 % in the purchase costs of the raw materials
were investigated. It appeared that the economic criteria of the extractive distillation
by gasoline are extremely sensitive. This causes a motivation to reject this option as
final design.
This implies that on basis of several criteria mentioned above, the normal distillation
followed by membrane purification appeared to be the most robust design. This
design has a Net Cash Flow of 272 kEURO, a Pay-Out Time of 3.4 years and a Rate
on Return of 29.4% before tax. The Discounted Cash Flow Rate on Return before tax
amounts to 24.6%.
i
Table of contents
1
2
3
4
5
6
7
Introduction
1
1.1
Motive
1
1.2
Purpose of the report
2
1.3
Description of report structure
2
Process options & selection
5
2.1
Process options
5
2.2
Criteria and selection
6
Basis of design
10
3.1
Description of the design
10
3.2
Process definition
10
3.3
Basic assumptions
12
3.4
Economic Margin
17
Thermodynamic properties
20
4.1
Operating window
20
4.2
Non-ideal equations for distillation
20
4.3
Choice of thermodynamic model
21
Process structure & description
30
5.1
Design criteria
30
5.2
Unit operations
30
5.3
Process chemicals
34
5.4
Utilities
35
5.5
Final process conditions
36
5.6
Process Flow Schemes
37
5.7
Process performance
39
Process control
42
6.1
General considerations
42
6.2
Control of a distillation column
42
6.3
Control of heaters and coolers
44
6.4
Control of a decanter
44
6.5
Control of the ultrafiltration unit in the extractive distillation by PAA
45
6.6
Control of the membrane unit in the normal distillation followed by membrane purification 45
Mass and heat balances
46
7.1
Mass and heat balances of the azeotropic distillation by toluene
46
7.2
Mass and heat balances of the extractive distillation by gasoline
47
ii
8
9
10
11
12
13
7.3
Mass and heat balances of the extractive distillation by PAA
48
7.4
Mass and heat balances of the normal distillation followed by membrane purification
48
Process and equipment design
50
8.1
Integration by process simulation
50
8.2
Equipment selection and design
51
8.3
Special issues
60
8.4
Equipment data sheets
61
Wastes
62
9.1
Identification of wastes
62
9.2
Biological treatment of waste water
63
9.3
Influence of process on wastes
64
Process safety
67
10.1
The Dow Fire and Explosion Index
67
10.2
Hazard and Operability Studies
69
Economy
71
11.1
Capital investment
71
11.2
Operating costs
73
11.3
Income
75
11.4
Cash flow
76
11.5
Economic criteria
77
11.6
Cost review
79
11.7
Sensitivities
79
11.8
Negative cash flows
81
Comparison and conclusions
82
12.1
Data validity
82
12.2
Purity and recovery
82
12.3
Process yields
83
12.4
Wastes
83
12.5
Process Safety
84
12.6
Economy
84
12.7
Selection for recommendation for further design
86
Recommendations
87
13.1
General recommendations
87
13.2
Azeotropic distillation by toluene
87
13.3
Extractive distillation by gasoline
88
13.4
Extractive distillation by PAA
88
iii
13.5
Normal distillation followed by membrane purification
88
Literature
90
Text symbols
92
Appendices:
Appendix 1
Appendix 2
Appendix 3
Appendix 4
Appendix 5
Appendix 6
Appendix 7
Appendix 8
Appendix 9
Appendix 10
Appendix 11
Appendix 12
Appendix 13
Appendix 14
Appendix 15
Appendix 16
Appendix 17
Appendix 18
Appendix 19
Appendix 20
Appendix 21
Appendix 22
Appendix 23
Pure component properties, toxicological and heat data, structure of
main components
Block schemes of the designs
Assignment
Comparison of the different processes
VLE data of regression for PAA design option
Utility costs
Utility summaries
Process flow schemes
Description of the Aspen Plus 10 files
Process stream summaries
Process yields
Heat and mass balances
Example calculations
Calculations of the columns, reboilers, condensers, coolers and heaters
Equipment summary & specification sheets of the azeotropic
distillation by toluene
Equipment summary & specification sheets of the extractive distillation
by gasoline
Equipment summary & specification sheets of the extractive distillation
by PAA
Equipment summary & specification sheets of the normal distillation
followed by membrane purification
Determination of the bundle diameter of the reboiler
Dow’s Fire and Explosion Index
Hazard and Operability studies
Investment and production costs
Column and tray layout
iv
1
1.1
Introduction
Motive
Ethanol is one of the basic components of the chemical industry. Ethanol is a clear,
colourless, flammable, oxygenated hydrocarbon with the chemical formula C2H5OH.
It is miscible in all proportions with water and also with ether, acetone, benzene, and
some other organic solvents. The binary mixture ethanol / water contains an
azeotrope. The azeotropic vapour-liquid equilibrium mixture, which occurs at 78.1C
and 1 bar, contains 95.57 w% ethanol and 4.43 w% water (ref. 37). The chemical
properties of ethanol are dominated by the functional - OH group, which can undergo
many industrially important chemical reactions, like, dehydration, halogenation,
estrification and oxidation.
For the production of ethanol various feedstocks and hence methods are used.
Fermentation alcohol can be produced from grain, molasses, fruit, wine, whey,
cellulose and numerous of other organic sources. More than 90 % of the world
production of ethanol is based on biological feedstocks. Synthetic alcohol may be
produced from crude oil, gas or coal, but plays a minor role in the world ethanol
production with a share of only 7 %.
The commonly known application of ethanol is within the field of alcoholic
beverages, but in the industry ethanol is also suitable for many other applications, like
industrial solvent and antiseptic. Ethanol is also used as raw material for the
preparation of many industrial organic chemicals, like acetaldehyde, butadiene,
diethyl ether, ethyl acetate, ethyl amines, ethylene, glycol ethers and vinegar.
Since 1970 a new application for ethanol has been found as fuel additive. In 1970 it
was realised that the petroleum stocks are limited and a search for alternative fuel
sources began. One of the alternatives can be found in gasohol, which is a fuel source
based on the use of ethanol obtained from natural sources. To produce gasohol an
extender is made from a mixture of gasoline (90 w%) and ethanol (10 w%). Gasohol
has a higher octane number and burns more slowly, coolly and completely than
gasoline, resulting in reduced emission of some pollutants. On the other hand ethanolbased gasohol is often expensive and energy intensive to produce. Nowadays the use
of ethanol as an additive to petrol is an important application, for example in Brazil.
Today, fuel ethanol accounts for roughly two thirds of the world ethanol production
(ref. 5).
This report covers the design of four ethanol dehydration processes, which produce
ethanol to be used as fuel additive. The (imaginary) contractor is an ethanol
fermentation plant in the republic of Lithuania. The fermented ethanol, which has a
concentration of 12 vol%, is not suitable for use as fuel additive until it is dehydrated
and purified till 99.8 vol%. To concentrate this polluted ethanol stream the
fermentation plant uses a distillation column to reach 88.8 vol% ethanol. The
pollutants in this stream are water, acetaldehyde, ethyl acetate, isobutyl alcohol and
isopentyl alcohol. This stream is the feed stream for the ethanol dehydration plant
discussed in this report. The pure component properties of all relevant substances are
listed in Appendix 1.
-1-
For ethanol dehydration various possibilities are feasible, but these vary tremendously
in costs. In the past many designs were based on energy consuming distillation towers
and used hazardous materials (like benzene, ref. 13) to achieve the required level of
dehydration. It is therefore a challenge to design a purification unit with modern
techniques, which has a low-energy consumption and uses less hazardous materials.
Furthermore due to the heavy competition in the field of ethanol production and the
small-scaled production level of the designed plant (see window on market situation)
it is necessary to produce at low costs. Therefore a comparison has to be made
between several possibilities. The ethanol dehydration possibilities considered in this
report are:
1. Azeotropic distillation by toluene, as reference case
2. Extractive distillation by gasoline
3. Extractive distillation by polyacrylic acid (PAA)
4. Normal distillation followed by membrane purification
The block schemes of these processes can be found in Appendix 2. The azeotropic
distillation by toluene is selected as reference case, because this process is based on
the well-known and commonly used plants. The extractive distillation by gasoline is
an elegant solution because no back-extraction of gasoline is needed for the
application of the ethanol as a fuel additive. The extractive distillation by PAA is
chosen, because it uses a relatively new entrainer. The distillation followed by
membrane purification is chosen because it is a completely different and promising
technique. In Chapter 2 the process selection is extendedly described.
1.2
Purpose of the report
The purpose of the report is to make a comparison between the four possibilities of
ethanol dehydration and to make a recommendation for a further design of one of the
process options. Three promising possibilities (2, 3 and 4) will be designed to a
conceptual state, based on steady state operation, and will be compared with a
traditional dehydration method (1) within the field of economy. The Dow’s Fire and
Explosion Index and a HAZOP study for critical pieces of equipment will be used to
compare the dehydration alternatives as far as safety is concerned. A review of the
waste water treatment will be used to compare the environmental impact of the
dehydration units.
On the basis of literature research the various possibilities of ethanol dehydration and
physical data of the components are investigated. The three possibilities and the
reference case are simulated in the computer program Aspen Plus 10, after checking
the theoretical data with the programmed data sets. Whenever possible, data from
literature is derived, but lacks of information are filled up by educated estimates. After
the equipment sizing a brief environmental and safety study this report has been
written.
1.3
Description of report structure
In this report the available process options and selection procedure is described in
Chapter 2. In Chapter 3 the basic assumptions and conditions of the design are made
-2-
clear. The thermodynamic properties and the chosen thermodynamic model are
defined in Chapter 4. Subsequently in Chapter 5 the process structures of the chosen
alternatives are extendedly treated. The equipment choice is explained and the streams
and utilities defined. The process control is described in Chapter 6, and the mass- and
heat balances are defined in Chapter 7. The design of the process equipment is made
explicit in Chapter 8. In this chapter the exact sizes of the equipment are described
and calculated. In Chapter 9 the wastes of the four process alternatives are mentioned,
subsequently in Chapter 10 the safety of the four options is investigated by a Fire &
Explosion Index and a Hazard and Operability study. In Chapter 11 all previous
subjects are used to obtain the economy. In Chapter 12 the conclusions are drawn.
Finally in Chapter 13 the recommendations are made.
In the frame below the world market situation for ethanol is dealt with. Also the
impact of the designed plant on this market is forecasted. Finally the patent situation
of several parts of the four design alternatives is mentioned.
Market situation for products and competitors1
The total world ethanol production in 1998 was approximately 31.2 billion litres. This was a
little downturn compared with previous years, due to a decrease of the Brazilian production
and the Asian financial crisis. The European ethanol production is approximate 15 % of the
world production and accounts 4.7 billion litres. Within Europe and especially within the
European Union there is a stimulation program to increase the use of ethanol as fuel additive
up from 5.6 % in 1997 to 12 % in 2010. Because the production of bio-fuels is more
expensive than conventional fuels, the manufacture of them will have to be subsidised. France
and Germany are the biggest producers of fuel ethanol in the European Union, with
respectively 500 million litres and 390 million litres. However, in Germany the ethanol
production is dominated by synthetic manufacturing, mainly by Hüls (177 million litres per
year) and Erdölchemie (75 million litres per year). The third largest ethanol producer in
Europe is the United Kingdom with a total production capacity of around 430 million litres
per year of which BP-Amoco alone accounts for 417 million litres divided over two sites.
This will change at the end of 2001, when BP-Amoco increases its production to 462 million
litres at the site of Grangemouth.
A relatively new fuel-ethanol project in the EU, which has come on stream around the turn of
the century, is Nedalco’s plant in Bergen op Zoom. This plant has a production capacity of 30
million litres per year. Other European ethanol production plants are Agroetanol’s facility in
Sweden (50 million litres per year) and the 100 million litres per year distillery in Cartagena,
Spain to be operated by Biocarburantes Espanoles.
In Eastern Europe the production of ethanol is dominated by the manufactures in the Russian
Federation. The total capacity in Russia can be estimated at 2.5 billion litres, with beverage
alcohol accounting for 60 %. The enormous capacity is hardly surprising, given the fact that
Russians drink almost 2.2 billion litres of pure ethanol per year. However due to the
dissolution of the former Soviet Union, restructuring and a huge illicit production, the Russian
government introduced a monopoly on the production of alcohol in 1997. After removal of
the state support the competitiveness of the Russian producers decreased and the import of
alcohol into the country has risen tremendously. Nowadays the illicit production of ethanol is
still increasing in the chaotic industry and trade policy of the Russian government.
The total EU and US ethanol exports to Eastern Europe in 1997 amounted 360 million litres,
of which over 270 million litres came from the US. The largest countries of destination in
1997 were Georgia (188 million), Ukraine (24 million) and Latvia (20 million).
1
This frame is based on ref 5.
-3-
It is believed that the political nature of fuel-ethanol production makes it unlikely that there
will be a consistently large international trade in this product. Fuel-ethanol programs have
been put in place to create additional demand for the purchase of feed stocks from domestic
farmers. It would run contrary to this intention if large-scale imports were allowed, as they
would support foreign farmers. A bio-fuel program usually incurs large costs, and it would
become completely unjustifiable if that money was spent on imports. As a result, world trade
will generally be limited to industrial and potable alcohol.
Review of joining the market
The designed ethanol dehydration plant will be connected with the fermentation plant and is
therefore fixed in size. In case of joining the market with the proposed ethanol productivity of
8.6 million litres per year, little impact will be exposed on the world market. The new plant
will only account for 0.2 ‰ of the world market and therefore has no influence at all at the
market price. The profitability of the new plant will be dependent on the current market price.
According to ref. 9 the price of ethanol is predicted to decrease over the next few years
because of better technologies and increasing capacity even though the price of the feedstocks
will increase. Additionally in many countries one or two companies control the production of
ethanol. These companies could provide rivals with a competitive edge. Therefore it could be
difficult to join the market with a relatively small plant.
Patent Situation
Patented processes, which will be redesigned for a specified feed stream, will increase the
total cost. Therefore the current patent situation is an important issue. Azeotropic distillation
by several hydrocarbon entrainers, like benzene, cyclohexane and toluene, is industrially
applied since 1903. As a result the azeotropic distillation by toluene is widely patented on unit
operation, process unit sequence and thermal integration. The use of gasoline in extractive
distillation to dehydrate aqueous ethanol was granted with a United States patent in 1952
(U.S. patent 2,591,672). Integrated distillation / membrane pervaporation plants are
commercially operated since the eighties. Both the design of hybrid systems and the use of
hydrophilic zeolite membranes are extensively patented worldwide. The use of polymeric
entrainers to break the azeotrope of ethanol / water systems is still in the experimental phase.
It is a relatively new area of investigation. Because of the experimental character of this
process option several possibilities still exist to patent the commercial operation of extractive
distillation by polyacrylic acid. So the use of PAA as an entrainer at plant scale is a feasible
process option concerning the patent situation. It is the only process option of which no
similar plants seems to exist yet.
-4-
2
Process options & selection
Designing requires a well-considered choice of the type of process, which can only be
made after a thorough investigation of the goal of the plant and the possible
alternatives to achieve this goal. Therefore the requirements and the process options
are described in the next paragraph.
2.1
Process options
As mentioned in Chapter 1 the ethanol / water feed stream is bought from a nearby
ethanol fermentation plant. The upstream section of this plant imposes the
specifications on the feed stream of the plant to be designed. Within the fermentation
plant the ethanol is upgraded from 12 vol% to 88.8 vol%. This ethanol needs to be
further purified in the designed dehydration unit up to 99.8 vol% for the use as fuel
additive.
To reach this requirement of purity, the azeotrope of the ethanol / water mixture has to
be broken. The alternatives to achieve the ethanol / water separation found in
literature are listed shortly in Table 2.1. The alternatives have been divided in four
categories, namely methods to reach the azeotrope, to break the azeotrope, methods
that can directly achieve the required purity and methods that achieve the required
purity by different techniques.
Table 2.1: List of separation alternatives for ethanol / water mixtures.
Type of
Ethanol
Estimate energy
Process
separation
w%
consumption
kJ/dm3ethanol
To azeotrope 89 – 96
Conventional distillation
2,600
To azeotrope 89 – 96
Multi-effect vacuum
2,000
To azeotrope 89 – 96
Vapour recompression
1,800
Azeotropic
96 – 100
Adsorptive dehydration by molecular
sieves
Adsorptive dehydration by solid agents
Adsorptive dehydration by zeolites
Azeotropic distillation by entrainers
Extraction by gasoline, hydrocarbons
Extraction by non-volatile components
Azeotropic
Azeotropic
Azeotropic
Azeotropic
Azeotropic
96 – 100
96 – 100
96 – 100
96 – 100
96 – 100
Azeotropic
Azeotropic
Azeotropic
96 – 100 Low pressure distillation (< 11,5 kPa)
96 – 100 Membrane Technology by pervaporation
96 – 100 Membrane Technology by vapour
permeation
-5-
1,500
500
1,500
2,600
2,200
2,000
Ref.
26
26
26
3,000
1,000
26, 1
26, 24
32, 19
26
8
37, 21,
18, 14
6
12, 17, 32
1,000
30
Table 2.1 (continued)
Complete
89 – 100
Complete
89 – 100
Complete
Complete
89 – 100
89 – 100
Adsorptive dehydration by solid agents
Convential ‘dual’ distillation with
entrainer
Extraction with supercritical carbon
dioxide
Solvent extraction
Vacuum distillation
Other
12-100
Membrane Technology
Complete
89 – 100
700
24
5,000
26, 21, 14
2,500
2,500
10,000
26, 37, 15
37
26
26, 37, 25
Each of these alternatives has advantages and disadvantages, so several criteria are
chosen to decide which alternatives will be designed in detail. Some alternatives can
be combined to use the advantages of these technologies. For example a hybrid
system of a distillation column followed by membrane purification. Another very
interesting option to purify the ethanol is the use of a hydrophobic zeolite membrane
(ref. 25). Such a membrane will let ethanol through as the permeate stream and will
let the water flow by. The inlet stream of the membrane is the aqueous ethanol stream
from the ethanol fermentation at 12 vol%. In this case no distillation columns are
necessary. This option is not chosen for the time being because it falls outside the
chosen specifications and battery limits (see Chapter 3, Basis of Design).
2.2
Criteria and selection
The criteria to make a selection between the process options are listed below.
-
Modern techniques
Most of the conventional processes are invented in the second half of the 20th
century. Therefore the conventional processes are already designed in detail and
optimised. The first literature on azeotropic distillation with benzene as entrainer
can be found in 1902 and is first patented in 1903. Almost a century of research
and optimisation has resulted in an enormous amount of capable entrainers and
dehydration methods, which are patented all. Therefore a challenge can be found
in using alternative techniques, which are only known for the last decades (see
window Chapter 1) and have potential to improve the ethanol dehydration, for
example in the field of economy, energy use or use of hazardous materials. The
most promising alternatives found in literature will be chosen.
-
Hazardous materials
In modern chemistry it is impossible to design and build a plant when the
environmental consequences are out of proportion. Within the process a minimum
of hazardous materials should be used. Whenever possible a hazardous material
should be avoided or replaced by a less hazardous material.
-
Economical profitability
The variable costs are mainly dependent on the energy consumption during the
operation of the plant, although some alternatives have relatively high fixed costs.
For the first rough distinguish in costs only the energy costs will take into account.
Estimates of the energy consumption can be found in Table 2.1. The energy
consumption of the dehydration methods are based on literature, but there is great
-6-
variety in methods used, extent of purification and year of design. The mentioned
number can only be seen as a rough estimate, because the proper costs of this
consumption will largely depend on the extent of optimisation and integration, the
costs of the equipment and entrainers used.
-
Number of options
The time limit of the CPD-project is set in advance to twelve weeks (480 hours
per person). Within this period only a certain amount of work can be done.
Therefore it is inevitable to set a maximum number of options to be worked out.
In our opinion the maximum feasible number of options to be designed is four,
within the imposed time limit. The number of free options is brought back to
three in practice by the requirement of a base reference case to compare the
designs mutually.
-
Data availability and data processing
To facilitate the process of designing a certain amount of data has to be available.
Literature and experts can for example provide this data. A second important
requirement is the possibility to apply the accumulated data in the processsimulating program Aspen Plus 10.
-
Personal interests
In the assignment a small number of alternatives is listed, which are preferred by
the (imaginary) contractor (Appendix 3). Also the interests of the company
Controlec is an important factor in the choice of alternatives. The students’ own
interests are also taken into account.
In Appendix 4 a comparison is made between the process options given in Table 2.1,
based on the criteria mentioned above. The chosen process options are:
1. Azeotropic distillation by toluene (standard case)
2. Extractive distillation by gasoline
3. Extractive distillation by a polymeric entrainer (polyacrylic acid)
4. Normal distillation followed by zeolite membrane purification
Each process is operated continuously, because this is the most economic and easiest
way of operation and all equipment used is capable of being operated continuously.
2.2.1
Azeotropic distillation by toluene
In azeotropic distillation a third component, in this case toluene, is added to the feed.
This component is called the entrainer or mass separating agent. This component
changes the vapour-liquid equilibrium of the ethanol-water mixture, by forming a new
azeotrope. This ternary azeotrope enables the recovery of pure components by using
three columns. The top stream of the first column is the ternary azeotrope, while the
bottom product contains a high concentrated ethanol mixture. This concentrated
mixture is separated from its impurities in the second column, where the 99.8 vol%
pure ethanol is recovered at the top and the impurities like ethyl acetate, and
acetaldehyde are the bottom product. The ternary azeotropic mixture of ethanol /
water / toluene and the redundant water are led to a decanter where the water phase is
separated. The azeotropic mixture is led to a third column where the toluene phase is
-7-
separated from its impurities and recycled to the first column. The block scheme is
represented in Appendix 2.
Benzene is the most common entrainer used in azeotropic distillation and is already
known for a long time. This entrainer however is not chosen because of its
carcinogenic properties. Other candidate third components are listed and treated in
Chapter 5. The well-known process with toluene as the entrainer will be used as a
reference.
2.2.2
Extractive distillation by gasoline
An extractive distillation is performed to separate ethanol-water mixtures by adding a
third component. This entrainer changes the vapour-liquid equilibrium of the original
mixture. When such a component is chosen, that breaks the azeotrope, it becomes
possible to dehydrate the ethanol mixture relatively easy. Extractive distillation can be
accomplished by using gasoline as an entrainer. Because the produced ethanol will be
used in gasohol, an inventive integration can be made between the ethanol
dehydration and a gasoline refinery. By using gasoline as entrainer, the ethanol-water
azeotrope will be broken. The ethanol-water separation can be fulfilled like an
ordinary extractive process, but large savings can be made. The gasoline entrainer is
part of the product, so it is not necessary to recover and recycle it to the distillation
process. The column will be operated in such way that the gasoline entrainer together
with the ethanol is removed as a bottom product, while the aqueous stream will be
recovered over the top. There is no need to separate the gasoline-ethanol mixture,
because gasohol is a mixture between gasoline and ethanol. This reduces the
investments costs and the variable production costs (energy costs) considerable. A
general process is presented in Appendix 2.
2.2.3
Extractive distillation by the polymeric entrainer polyacrylic acid (PAA)
Another possibility to accomplish an extractive distillation process is the use of the
polymer entrainer PAA. This entrainer changes the vapour liquid equilibrium by
breaking the ethanol / water azeotrope. By adding PAA in the distillation column the
pure ethanol can be obtained overhead. The water-entrainer mixture is fed to an
ultrafiltration unit, where the water is separated from the entrainer. The entrainer,
dissolved in a small amount of water, is recycled to the distillation column. The
general process is represented in Appendix 2.
2.2.4
Normal distillation followed by membrane purification
Membranes can separate water from water-alcohol mixtures in a much more
economical way than by conventional distillation. These membranes can overcome
the azeotropic barrier and so can obtain the required purity of ethanol. Membranes are
especially useful to separate mixtures near the azeotropic composition. The
advantages of membrane technology are the low operation costs and the breaking of
the azeotrope without the aid of a solvent. A disadvantage is the low flux through the
membrane, which illustrates the need for multiple membrane-sections in series. A
-8-
rather new technology is the technology of zeolite membranes. These membranes can
selectively remove one component. A big advantage of zeolite membranes is that they
have a larger flux ( 4 kg/(m2.h)) through the membrane than polymeric membranes
( 0.2 kg/(m2.h)). Therefore less units will be needed when zeolite membranes are
used. Because of these advantages zeolite membranes will be used in the design,
which is shown in Appendix 2. A normal distillation column will purify the feed
stream to or near the azeotrope. The azeotrope comes overhead and water flows over
the bottom. Subsequently the top stream will be purified using the zeolite membranes.
The membrane sections separate the water by pervaporation. Another option is the
vapour permeation technology.
-9-
3
Basis of design
3.1
Description of the design
Pure ethanol has a wide range of applications in both the chemical and the consumer
industry. In the last decades ethanol is discovered as a valuable component in fuel.
Since then ethanol mixed with gasoline to produce gasohol. Before ethanol can be
used for most purposes the raw, aqueous ethanol needs to be purified from water. This
imposes a difficulty because the mixture of ethanol and water contains an azeotrope.
This binary azeotrope of the ethanol / water system is situated at 95.57 w% ethanol
and 4.43 w% water at 78.1C and 1 bar. This means that purification by normal
distillation cannot recover pure ethanol. To acquire the pure ethanol the azeotrope in
the ethanol / water system has to be broken. Various dehydration processes can
achieve the desired separation (see Chapter 2).
For the production of dehydrated ethanol, which is used as an additive to gasoline, a
purity of 99.8 vol% is required. In the upsteam production process ethanol is produced
in a stream of about 12 vol% aqueous ethanol. To concentrate this stream a distillation
column is used to reach 88.8 vol% ethanol. The ethanol stream leaving this distillation
column is the feed stream of the design. In the ethanol feed stream there are some
impurities as acetaldehyde, ethyl acetate, isobutyl alcohol and isopentyl alcohol
present in relatively small amounts, but there are no requirements to keep this
impurities out of the ethanol stream (see Appendix 3).
3.2
Process definition
A well-considered choice in the type of process can only be made after a thorough
investigation of the goal of the plant and the possible alternatives to achieve this goal.
3.2.1
Process concepts chosen
To reach the requirements of purity, the azeotrope of the ethanol / water mixture has
to be broken. There are many alternatives available to achieve the desired ethanol /
water separation (see Chapter 2, Table 2.1). All these alternatives have advantages
and disadvantages. To make a choice between the suitable process options a selection
is made between them on the basis of several criteria. The following criteria are used
to decide which alternatives will be designed in detail (see Appendix 4):
- Modern techniques
- Hazardous materials
- Economical profitability
- Number of options
- Data availability
- Personal interests
After a comparison between the process options, based on the criteria mentioned
above. The chosen process options are:
-10-
-
Azeotropic distillation by toluene
Extractive distillation by gasoline
Extractive distillation by polyacrylic acid (PAA)
Normal distillation followed by membrane purification
Each process is operated in a continuous mode.
The option extractive distillation by gasoline is directly producing gasohol in contrast
to the other three options that produce pure ethanol. Because the specification of 99.8
vol% pure ethanol does not hold anymore, the following specifications are imposed.
The gasohol to be produced must contain 10 w% ethanol. This 10 w% of ethanol in
gasohol has to be 99.8 w% pure. This imposes a maximum allowable amount of water
in gasohol of 0.2 w% of the ethanol present in gasohol.
3.2.2
Block schemes
The block schemes for the four process options are provided in Appendix 2. Only the
significant pieces of equipment are represented in this block scheme. In the block
schemes the total mass streams (ton/annum) and yields (ton/ton product) are given. In
the separate blocks the process conditions are displayed.
3.2.3
Thermodynamic properties
To calculate the exact thermodynamic properties the liquid activity coefficients and
vapour fugacities are necessary (See Chapter 4). For an estimation of these parameters
several methods are in use. To determine the right thermodynamic method, the
feasible models are investigated in Aspen Plus 10. These models are the Wilson, the
NRTL, the UNIQUAC and the UNIFAC model. All these models are applicable to
(highly) non-ideal mixtures, Vapour-Liquid Equilibria (VLE) and except for the
Wilson model they are also applicable to two liquid phases.
For the separation of ethanol and water the models NRTL and UNIQUAC are most
likely to be used. When the separation is accomplished with the aid of a mass
separating agent, for example toluene, the system changes from two phases to three
phases. Also in this case the NRTL and the UNIQUAC models are valid.
To determine the influence of the model both the NRTL and the UNIQUAC method
are evaluated by creating x,y-diagrams and residue curves for the binary systems and
the ternary system of ethanol, water and toluene (see Chapter 4, Figure 4.1 and 4.2).
The differences between the available models mutually and with literature are small.
Nevertheless a choice has to be made between the UNIQUAC and the NRTL model.
Both models are appropriate, but here the choice is made for the UNIQUAC model,
because of its wide acceptance in literature and its accuracy in representing VLE data
for a wide range of systems.
Because the thermodynamic model has to be useable for the systems with an entrainer
as well, the validity of the UNIQUAC model for this systems is investigated (See
Chapter 4). It appears that the UNIQUAC model can also be used as thermodynamic
-11-
model for the azeotropic distillation with toluene as mass separating agent and for the
extractive distillation with respectively gasoline and polyacrylic acid.
3.2.4
Pure Component Properties
The list with pure component properties, toxicological and heat data are provided in
Appendix 1.
3.3
3.3.1
Basic assumptions
Plant capacity
The ethanol feed stream of the plant is fixed on 10,054 ton/year with 88.8 vol% (86.3
w%) ethanol (see Appendix 3). The goal of the recovery of ethanol from the feed
stream is set at a minimum of 99 w%. The actually attained recovery and purity of the
ethanol product is calculated for each process option in Chapter 5. The plant is
operated for 8,620 stream hours per year. The plant capacity of the four dehydration
units is summarised in Table 3.1.
Table 3.1: Annual plant capacity for the four process options.
Process option
Feed capacity (t/a)
Azeotropic distillation by toluene
10,054
Extractive distillation by gasoline
10,054
Extractive distillation by PAA
10,054
Distillation followed by membrane
purification
10,054
Production capacity (t/a)
8,286
86,737*
8,689
8,599
* Production of gasohol
The extractive distillation by gasoline has a higher production capacity than the other
options, because in gasohol the ethanol is mixed with a large amount of gasoline.
Once every five year a major maintenance will be necessary. Because the production
equipment also needs this maintenance, this will not be taken into account. The
repairs, on the other hand, do have to be taken into account (Appendix 3). The time
for repairs is set at 150 hours. To storage the feed during this repairs there has to be a
storage tank on the site. Because the repairs have to be fulfilled within 150 hours, the
storage tank capacity has to be at least 225 m3. The feed storage tank is assumed to be
on the site. The economical plant life is assumed to be 15 years.
3.3.2
Location
The ethanol dehydration plant will be located next to the ethanol fermentation site.
The plant is located is in the Republic of Lithuania. Because other chemical industry
surrounds the dehydration unit, it is assumed that excellent utilities are available
(Appendix 3). These utilities are listed in Table 3.2.
-12-
Table 3.2: Available utilities at the plant site.
Utility
Medium pressure steam at 10 bar
Electric power
Cooling water
Nitrogen
Price
12.5
50
450
1.35
(Euro/ton)
(Euro/MWh)
(Euro/kton)
(Euro/ton)
Also other necessary material, including chemicals, are assumed to be available at
market prices. The infrastructure is fully developed and electricity, air and sewerage
facilities can be easily constructed.
3.3.3
Battery limits
The ethanol is produced by fermentation in the nearby production plant. During the
ethanol production some side reactions take place, so small amounts of aldehydes,
higher alcohols and esters are formed. These components will be converted
respectively to acetaldehyde, ethyl acetate, isobutyl alcohol and isopentyl alcohol. At
the nearby production site a predistillation already takes place to concentrate the
ethanol to a value of 88.8 vol%. The preconcentrated ethanol is stored in a tank at 30
C and 1 bar. From this storage tank the feed stream will cross the battery limit. Inside
the battery limits the ethanol is dehydrated by the chosen methods. In each design the
dehydration is mainly performed by distillation columns (see Appendix 2). Out of the
battery limit flows respectively a dehydrated ethanol stream or gasohol stream.
Furthermore a waste water stream comes out the dehydration plant crossing the
battery limits. The treatment of the waste water will not be included in this design, but
the cost of the treatment on the other hand must be taken into account. In the case of
the azeotropic distillation by toluene (30 ºC, 1 bar) and extractive distillations by
gasoline (20 ºC, 1 bar) and PAA a stream of entrainer flows into the battery limit from
a storage tank. An amount of the entrainer flows out of the system as impurities in the
ethanol or waste water stream.
3.3.4
Definition in- and outgoing streams
Various streams enter the battery limits of the four options. In each design an ethanol
feed stream is present and enters the battery limit. Besides this feed stream an
entrainer is added in the azeotropic distillation by toluene and in the extractive
distillation by gasoline, respectively a small quantity of toluene and a large quantity of
gasoline. These streams also pass the battery limits. The specifications of the ethanol
feed stream are listed in Table 3.3. In the design all components of the ethanol feed
stream have the maximal allowable concentration: the worst case design.
-13-
Table 3.3: Conditions of the ethanol feed stream for the four alternatives.
Stream Name :
Ethanol feed
Component
Units
Specification
Additional Information
Available Design Notes
Ethanol
vol%
88.8
(1)
 88.8
Acetaldehyde
mg/m3
300
(1)
 300
Isobutyl alcohol
mg/m3
1250
(1)
 1250
Isopentyl alcohol
mg/m3
3750
(1)
 3750
Ethyl acetate
mg/m3
500
(1) (1) As ‘worst case’ scenario.
 500
Process Conditions and Price
Temperature
30
C
Pressure
Bar
1
Phase
V/L/S
L
Price ethanol
EUR/ton
250
A representative composition for the gasoline entrainer is listed in Table 3.4.
Table 3.4: Composition of the gasoline feed stream.
Stream Name :
Gasoline feed
Component
Units
Toluene
w%
1-Hexene
w%
2-Methyl-2-butene
w%
Methyl cyclopentane
w%
Methyl cyclohexane
w%
n-Pentane
w%
2-Methylbutane
w%
n-Hexane
w%
2-Methylpentane
w%
3-Methylpentane
w%
n-Heptane
w%
Process Conditions and Price
Temperature
C
Pressure
bar
Phase
V/L/S
Price
EUR/ton
Composition
10
6
4
13
7
15
8
12
12
6
7
20
1
L
1,200
The in- and outgoing streams of feedstocks, products and wastes crossing the battery
limits are summarised for each process in Table 3.5 till Table 3.8. The only
specification for the ethanol product stream given, is that it has to contain at least 99.8
vol% ethanol. In all four designs there is a waste water stream present.
-14-
Table 3.5: Streams passing battery limits in the azeotropic distillation by toluene.
Stream name:
Ethanol feed Toluene feed
Ethanol
Waste water
product
Component
Mw
kg/s
kg/s
kg/s
kg/s
Ethanol
46.07
0.280
0.267
0.013
Water
18.02
0.042
0.000
0.042
Toluene
92.14
0.003
0.000
0.003
Isopentyl alc.
88.15
0.002
0.000
0.002
Isobutyl alc.
74.12
0.001
0.000
0.000
Ethylacetate
88.11
0.000
0.000
0.000
Acetaldehyde
44.05
0.000
0.000
0.000
Total
0.324
0.003
0.267
0.060
Enthalpy
kW
-2,356
0
-1,607
-744
Phase
L/V/S
L
L
L
L
Pressure
bar
1.0
1.0
1.0
1.0
Temperature ºC
30.0
30.0
30.0
40.0
Price
EUR/ton
250
450
550
-650
Table 3.6: Streams passing battery limits in the extractive distillation by gasoline.
Stream name:
Gasoline feed Ethanol feed
Water
discharge
Component
Mw
kg/s
kg/s
kg/s
Ethanol
46.07
0.000
0.280
Water
18.02
0.000
0.042
0.042
Toluene
92.14
0.251
0.000
Ethylacetate
88.11
0.000
0.000
Isobutyl alcohol
74.12
0.000
0.001
Isopentyl alcohol
88.15
0.000
0.002
Acetaldehyde
44.05
0.000
0.000
1-Hexene
84.16
0.151
2-Methyl-2-butene
70.14
0.101
Methylcyclopentane 84.16
0.327
Methylcyclohexane 98.19
0.176
N-pentane
72.15
0.377
N-hexane
86.18
0.302
2-Methylpentane
86.18
0.302
3-Methylpentane
86.18
0.151
N-Heptane
100.25
0.176
2-Methylbutane
72.15
0.201
Total
2.513
0.324
0.042
Enthalpy
kW
-4,674
-2,356
-660
Phase
L/V/S
L
L
L
Pressure
bar
1.0
1.0
1
Temperature
ºC
20.0
30.0
37.9
Price
EUR/ton
1,200
250
-
-15-
Gasohol
kg/s
0.280
0.000
0.251
0.000
0.001
0.002
0.000
0.151
0.101
0.327
0.176
0.377
0.302
0.302
0.151
0.176
0.201
2.795
-6,318
L
1.0
30.0
1,296
Table 3.7: Streams passing battery limits in the extractive distillation by PAA.
Stream name:
Ethanol feed
Ethanol product
Waste water
Component
Mw
kg/s
kg/s
kg/s
Ethanol
46.07
0.280
0.279
0.000
Ethyl acetate
88.11
0.000
0.000
0.000
Acetaldehyde
44.05
0.000
0.000
0.000
Water
18.02
0.042
0.000
0.042
Isobutyl alcohol
74.12
0.001
0.000
0.001
Isopentyl alcohol
88.15
0.002
0.000
0.002
PAA
2000
0.000
0.000
Total
0.324
0.280
0.044
Enthalpy
kW
-2,356
-1,680
-672.2
Phase
L/V/S
L
L
L
Pressure
bar
1.0
1.0
1.0
Temperature
ºC
30.0
30.0
40.0
Price
EUR/ton
250
550
-125
Table 3.8: Streams passing battery limits in the normal distillation followed by membrane
purification.
Stream name
:
Ethanol feed
Ethanol product
Waste water
Component
Mw
kg/s
kg/s
kg/s
Ethanol
46.07
0.280
0.277
0.003
Water
18.02
0.042
0.000
0.042
Isopentyl alcohol
88.15
0.002
0.000
0.002
Isobutyl alcohol
74.12
0.001
0.000
0.001
Ethylacetate
88.11
0.000
0.000
0.000
Acetaldehyde
44.05
0.000
0.000
0.000
Total
0.324
0.277
0.047
Enthalpy
kW
-2,356
-1,666
-686.2
Phase
L/V/S
L
L
L
Press.
bar
1.0
1.0
1
Temp
ºC
30.0
30.0
40
Price
EUR/ton
250
550
-224
3.3.5
General assumptions
In designing the four process alternatives several general assumptions have been
made. Further design choices are dealt with throughout the report and are not listed
here.
The general assumptions made are:
1. The feed stream of 88.8 vol% ethanol has the conditions of T = 30 °C and p = 1
bar. These conditions are assumed, because no conditions are given in the
assignment (Appendix 3).
2. The ethanol product stream that crosses the battery limits has the conditions T =
30 °C and p = 1 bar. So the specification for the product purity (99.8 vol%) holds
for these conditions. (Obviously this assumption doesn’t hold for the gasohol
production plant.)
-16-
3. The mass flows of all components in the feed and product stream are calculated by
assuming that these streams are ideal liquids. The mass or volume flows are
calculated using the densities of each component separately at the stated
conditions, so not one density of the whole mixture (non-ideal). To illustrate this
Table 3.9 is added. The error made by assuming this is very small and therefore
difficult calculations are avoided.
4. The cooling water is available at 20 ºC and 3 bar and it is allowed to be disposed
off at 40 ºC after using as utility.
5. Pressure loss due to pipeline friction is neglected.
6. For condensers, reboilers, heaters and coolers a heat loss of 5 % is taken into
account. From there on the needed exchange area and utility amount is calculated.
7. All pumps have an outlet flow at 0.50 m above ground level. This is taken into
account to calculate the duties of the pumps that pump liquid up to a certain level.
8. The disposal of the condensed utility steam is not taken into account.
Table 3.9: The composition of the feed stream.
Name :
Ethanol feed
Component
MW
kg/h
kmol/h
Ethanol
46.07 1,006.2
21.841
Water
18.02 150.95
8.379
Ethyl acetate
88.11
0.725
0.008
Acetaldehyde 44.05
0.435
0.010
Isobutyl
74.12
1.813
0.024
alcohol
Isopentyl
88.15
5.438
0.062
alcohol
Total
1,165.56
30.324
Phase
L/V/S
L
Pressure
bar
1.0
Temperature ºC
30.0
m3/h
1.288
0.152
0.001
0.001
0.002
w%
0.863
0.130
0.001
0.000
0.002
mol%
0.720
0.276
0.000
0.000
0.001
vol% kg/m3(ref.28)
0.888
781.5
0.105
993.7
0.001
887.7
0.000
767.6
0.002
790.4
0.007
0.005
0.002
0.005
1.450
1.000
1.000
1.000
803.2
The density of isobutyl alcohol is estimated using the values given underneath: B*C/A
Density of 2-methyl-2-propanol at 20 ºC (kg/m3),
A:
788.8 (ref. 28)
3
Density of 2-methyl-2-propanol at 30 ºC (kg/m ),
B:
777.6 (ref. 28)
3
Density of isobutyl alcohol at 20 ºC (kg/m ),
C:
801.8 (ref. 20)
3.4
Economic Margin
To determine the maximum allowable investments for the designs the economic
margin is calculated. This margin is the difference between income from sales minus
the costs of the feedstock. According to ref. 10, the market price for ethanol fuel grade
is 550 EUR/ton. The market price of ethanol 88.8 vol% is not available, so the price
has to be assumed. As the ethanol prices are not proportional to the percentage
ethanol, because of the efforts to overwin the azeotrope, a price of 88.8 vol% ethanol
is assumed at 45 % of the fuel grade price. So the market price of raw ethanol is set at
250 EUR/ton. The prices for all the feedstocks, entrainers and products are tabulated
in Table 3.10.
-17-
Table 3.10: Costs raw materials and incomes from product sales.
Raw materials:
Unity
1
Ethanol 88.8 w%
ton
Gasoline2
ton
Toluene2
ton
1
Polyacrylic acid
ton
EUR per unity
250
1,200
450
1,300
Products:
Ethanol
Gasohol
EUR per unity
550
1,296
Unity
ton
ton
1 Estimation
2 ref. 10
The economic margin per year can be defined as:
Margin = Total Product Revenues - Total Feedstock Costs
(3.1)
The margin for each process option is summarised in Table 3.11.
Table 3.11: Economic Margin.
Process option
Production
capacity
(t/a)
Azeotropic distillation
by toluene
8,286
Extractive distillation
by gasoline
86,737
Extractive distillation
by PAA
8,689
Normal distillation
followed by
membrane purification
8,599
Product
revenues
(kEUR/a)
Feedstock costs
(kEUR/a)
Margin
(kEUR/a)
4,565
2,562
2,003
112,428
96,113
16,315
4,616
2,513
2,103
4,406
2,513
1,893
From the table it can be seen that the option extractive distillation by gasoline has the
largest margin due to the high production capacity. The other three options do not
vary much from each other. With the economic margin the maximum allowed
investment is calculated by a discount cash-flow analysis as described in ref. 33,
p.239. This method is used to calculate the present worth of future earnings. This can
be used to determine the maximum allowed investment (equation 3.2). A discounted
cash-flow rate of return (r’) of 10 % is assumed and a plant life (n) of 15 years.
n t

n 1
Margin
1  r 
' n
 Maximum Allowable Investment
(3.2)
The maximum allowable investments calculated from equation 3.1 for each process
option are summarised in Table 3.12.
-18-
Table 3.12: Maximum allowable investment for the four designs
Process option
Maximum allowable investment (kEUR)
Azeotropic distillation by toluene
17,238
Extractive distillation by gasoline
140,408
Extractive distillation by PAA
18,099
Distillation followed by membrane
purification
16,291
The results will be compared with the economic evaluation in Chapter 11.1.
-19-
4
Thermodynamic properties
4.1
Operating window
The thermodynamic relations and estimation methods used in the process designs
should be valid for the temperatures and pressures occurring in the equipment.
Therefore an operating window is defined for each process and is shown in Table 4.1.
Table 4.1: Operating window for the alternatives.
Process
Temperature range (ºC)
Azeotropic distillation by toluene
30.0 – 112.7
Extractive distillation by gasoline
20.0 – 81.6
Extractive distillation by PAA
30.0 – 92.9
Normal distillation followed by
30.0 – 94.8 (distillation)
membrane purification
3.8 – 120.0 (membrane)
4.2
Pressure range (bar)
1 – 1.5
1 – 2.4
1 – 3.0
1 – 1.2 (distillation)
0.0 – 4.3 (membrane)
Non-ideal equations for distillation
The main separation of the water from the ethanol takes place in the distillation
columns in each design. On each stage in the distillation column both vapour and
liquid phases are present and are in (thermodynamic) equilibrium. The equilibrium of
component j in the vapour and liquid phase is based on the equality of the fugacity fˆ
in both phases (ref. 34, p.338):
fˆjl  fˆjv
in which:
fˆji
(4.1)
Fugacity of mixture of component j in phase i
This equation (4.1) accounts under the restrictions of constant temperature and
pressure. For component j in the non-ideal vapour phase the fugacity and the fugacity
coefficient are related as follow (ref. 34, p.366):
fˆjv  y j  ˆjv  P
(4.2)
in which:
yj
ˆv
Mole fraction of component j in the vapour phase
P
Total pressure
j
Fugacity coefficient of the vapour phase mixture
For component j in the non-ideal liquid phase a similar relation between fugacity and the
activity coefficient  exists (ref. 34, p.368):
fˆjl  x j   j  f jl
(4.3)
-20-
in which:
xj
Mole fraction of component j in liquid phase
j
Activity coefficient of component j
i
j
Fugacity of component j in phase i.
f
Using equations (4.1) to (4.3) the vapour-liquid equilibrium constant K for component j is:
Kj 
yj
xj

 j  f jl
(4.4)
ˆjv  P
This quantity determines the relative volatility of two components and thus the
separation of these components.
Other thermodynamic properties like heat capacities and enthalpies for the pure
components can be found in Appendix 1.
4.3
Choice of thermodynamic model
As can be seen in equations (4.1) till (4.4) the liquid activity coefficients and vapour
fugacities are necessary to calculate the exact thermodynamic properties. For an
estimation of these parameters several methods are in use. To determine the right
thermodynamic model, the feasible models are investigated in Aspen Plus 10. These
models are the Wilson, the NRTL, the UNIQUAC and the UNIFAC model. All these
models are applicable to (highly) non-ideal mixtures, vapour-liquid equilibria (VLE)
and except for the Wilson model they are also applicable to two liquid phases.
Although a phase separation into two liquids between ethanol and water is not
expected, but is expected between the water or ethanol and one or more entrainers, the
Wilson method seems to be more inaccurate in this design than the other methods.
The UNIFAC (UNIQUAC functional group activity coefficient) model is valid in the
temperature range from 2 ºC to 202 ºC and in a pressure range from 0 to 4 bar. This
model is an extended form of the UNIQUAC method and applies different models to
estimate unknown thermodynamics properties from the group contributions instead of
molecular contributions. A disadvantage of this method is the inaccurate values of
parameters. Because this model estimates a large number of parameters it is only used
in cases where almost all parameters are unknown. This is not the case with ethanol /
water mixtures, so the UNIFAC method will not be used.
For the separation of ethanol and water the models NRTL and UNIQUAC are most
likely to be used. When the separation is accomplished with the aid of a mass
separating agent, for example toluene, the system changes from two phases (vapour
liquid) into three phases (vapour liquid liquid). Also in this case both the models can
be used. To determine the influence of the model both the NRTL and the UNIQUAC
method are evaluated by creating x,y-diagrams and residue curves for the binary
systems and the ternary system of ethanol, water and toluene (see figures below). In
these diagrams the azeotropic points should be seen at or near the value in literature.
In Table 4.2 all relevant azeotropic points found in ref. 20 (p. 6-221, 6-239) are listed.
-21-
Table 4.2: Azeotropic points at atmospheric pressure according to ref. 20.
mixture
phases
Wtfrac (%)
Molfrac (%)
Ethanol / water
LV
96 / 4.0
90 / 10
Ethanol / toluene
LV
68 / 32
81 / 19
Water / toluene
L 1 L2 V *
13.5 / 86.5
44.4 / 55.6
Ethanol / water /
toluene
L1 L2 V *
37 / 12 / 51
40 / 33 / 27
T (ºC)
78.17
76.7
84.1
74.4
* : The azeotropic mixture is the vapour phase. The two liquid phases are almost immiscible.
4.3.1 Ethanol / water mixture
The most important azeotrope is that of ethanol / water. For this mixture the VLEcurve obtained from the flowsheet calculation program Aspen Plus10 is compared
with the theoretical data of Perry (ref. 29, p.13-12) in Figures 4.1 and 4.2.
1
0.9
vapour molefraction ethanol (-)
0.8
0.7
0.6
Uniquac 1 bar
0.5
NRTL 1 bar
Perry
0.4
0.3
0.2
0.1
0
0
0.1
0.2
0.3
0.4
0.5
0.6
0.7
0.8
0.9
1
liquid molefraction ethanol (-)
Figure 4.1: Comparison of the x,y-diagram of the ethanol / water system from Aspen Plus 10
models and Perry.
-22-
1
0.98
Vapour fraction ethanol (-)
0.96
0.94
0.92
Uniquac
0.9
NRTL
Perry
0.88
0.86
0.84
0.82
0.8
0.8
0.82
0.84
0.86
0.88
0.9
0.92
0.94
0.96
0.98
1
liquid molefraction ethanol (-)
Figure 4.2: Zoom-in of the azeotropic point in the mixture ethanol-water.
As can be seen from Figure 4.2 the Aspen models are consistent with the literature
values given in Table 4.2. Because both the liquid and the vapour phase of the system
are highly non-ideal the NRTL-HOC and the UNIQ-HOC method are also evaluated.
These models use the Hayden-O’Connell equation to describe the non-ideal vapour
phase. The plot is shown in Figure 4.3.
-23-
1
0.9
vapour mole fraction ethanol (-)
0.8
0.7
0.6
UniQuac
NRTL
0.5
UniQ-Hoc
NRTL-Hoc
0.4
0.3
0.2
0.1
0
0
0.1
0.2
0.3
0.4
0.5
0.6
0.7
0.8
0.9
1
liquid mole fraction ethanol (-)
Figure 4.3: Comparison thermodynamic models UNIQUAC, UNIQUAC-HOC, NRTL and
NRTL-HOC.
As can be seen in Figure 4.3, the difference between the available models is small.
This is probably the result of the wide acquaintance of the used substances. Therefore
the UNIQUAC-model and the NRTL-model probably will take some non-ideal
behaviour of the vapour phase into account. The two HOC-models are especially
suitable for vapour reaction. Because no such reaction takes place, a choice has to be
made between the UNIQUAC and the NRTL model. Both models are appropriate, but
the UNIQUAC model is used in this situation, because of its wide acceptance in the
literature and its accuracy in representing VLE data for a wide range of systems (ref.
2). A small disadvantage of the UNIQUAC (but also of NRTL) method is that the
parameters all inherit a Boltzmann-type T dependence from the origins of the
expressions for GE, but it is only approximate. (ref. 29, p. 4-23)
4.3.2 Other components
Within the four design options three options make use of a third component to
separate the water from the ethanol, namely the entrainers toluene, polyacrylic acid
and gasoline. For this third component the thermodynamic model should also be
valid. The UNIQUAC model accounts for liquid-vapour equilibria as well as liquidliquid-vapour equilibria. Therefore the UNIQUAC model can be used in all four
alternatives. But in each design option the validity of UNIQUAC has to be checked:
in Aspen Plus 10 for the interaction between ethanol, water and the third component.
-24-
Azeotropic distillation by toluene
To validate the use of UNIQUAC for the azeotropic distillation by toluene, the
thermodynamic properties of ethanol, water and toluene are investigated. This is done
by comparing the equilibrium-curves of the binary mixtures and in a ternary mixture
with literature of Table 4.2 above.
Table 4.3: Azeotropic points at atmospheric pressure according to UNIQUAC-Aspen Plus 10.
mixture
Phases
molfrac (%)
Ethanol / water
LV
90 / 10
Ethanol / toluene
LV
81 / 19
Water / toluene
L 1 L2 V *
56 / 44
Ethanol / water / toluene
L1 L2 V *
46 / 28 / 26
* : The azeotropic mixture is the vapour phase. The two liquid phases are almost immiscible.
1
Aspen-UNIQUAC simulations are done to analyse all the equilibria of the mixtures of
ethanol, water and toluene. In Table 4.3 all azeotropic points found are summarised.
First the equilibrium of ethanol and toluene is simulated. This simulation gives an
azeotrope of about 81 mol% ethanol, showed in Figure 4.4. Table 4.2 gives a
literature value for the azeotrope of 81 mol% ethanol. So the UNIQUAC result
corresponds well with literature.
Y-x for ETHANOL/TOLUENE
0.2
Vapor Molefrac ETHANOL
0.4
0.6
0.8
1.0133E+05 N/sqm
0
0.05 0.1 0.15 0.2 0.25 0.3 0.35 0.4 0.45 0.5 0.55 0.6 0.65 0.7 0.75 0.8 0.85 0.9 0.95
Liquid Molefrac ETHANOL
Figure 4.4: X,y-diagram of the ethanol / toluene system from Aspen-UNIQUAC.
-25-
1
1
Y-x for WATER/TOLUENE
0.2
Vapor Molefrac WATER
0.4
0.6
0.8
1.0133E+05 N/sqm
0
0.05 0.1 0.15 0.2 0.25 0.3 0.35 0.4 0.45 0.5 0.55 0.6 0.65 0.7 0.75 0.8 0.85 0.9 0.95
Liquid Molefrac WATER
Figure 4.5: X,y-diagram of the water / toluene system from Aspen-UNIQUAC.
As mentioned in Table 4.2 the mixture of water and toluene is a three phase mixture:
two liquid phases and a vapour phase. As can be seen from Figure 4.5 the two liquid
phases are nearly pure water and toluene. This is because the inorganic and organic
components are not miscible. The vapour phase contains the azeotropic mixture.
Figure 4.5 shows that this azeotrope lies at 55.8 mol% water. In ref. 11 a value of 55.5
mol% water is given. A remarkable difference is seen between these two similar
values and the literature value given in Table 4.2, 55.6 mol% toluene. Other Aspen
models like NRTL and UNIFAC were consulted and all give about the same results as
UNIQUAC. Although it seemed very remarkable that literature can not reach
accordance, it is assumed that (ref. 11) and the several Aspen-simulations are correct.
Finally a residue curve is made for the ternary mixture of ethanol, water and toluene,
as can be seen in Figure 4.6. The literature value stated in Table 4.2 is 40 mol%
ethanol, 33 mol% water and 27 mol% toluene. The residue curve shown in Figure 4.6
gives an azeotropic point at about 46 mol% ethanol, 28 mol% water and 26 mol%
toluene.
This result is quite reasonable. According to the residue curve a greater fraction
ethanol will go overhead with the azeotropic mixture in a distillation column. This
implies that a column, where ethanol is supposed to be received as the bottom stream,
performs better in reality than a distillation column designed in a flowsheet program
with the UNIQUAC parameters.
-26-
1
Residue curve for ETHANOL/WATER/TOLUENE
0.2
0.8
ER
AT 0.4
Mo
l
0.6 efrac
W
TO
rac
lef
Mo6
0.
LU
EN
0.4 E
0.2
0.8
0.2
0.4
0.6
Molefrac ETHANOL
0.8
Figure 4.6: The residue curve for ethanol, water and toluene.
Extractive distillation by gasoline
For the separation of ethanol and water gasoline is used as an entrainer. Some
components of gasoline, like n-heptane, form an azeotrope with ethanol. Gasoline and
water are immiscible components and form two liquid phases. Using Aspen Plus 10
these effects come back in the UNIQUAC model, so this model can be used for the
designing.
Extractive distillation by PAA
The extractive distillation by PAA is a special case. In literature (ref. 2) the separation
of an ethanol / water mixture is mentioned due to the shift of the azeotropic point by
adding the polymer polyacrylic acid. In this article experimental values for ethanol /
water / PAA phase-equilibria are given. The authors found that the ethanol / water
azeotrope will disappear when 0.45 w% PAA is added to the mixture. Because the
available Aspen Plus 10 is not capable of simulating polymers and therefore is unable
to shift the azeotrope, the VLE data from literature are fit to the UNIQUAC model.
The regression results are shown in Figure 4.7, and the VLE data are listed in
Appendix 5.
Using the maximum-likelihood method the UNIQUAC parameters, according to
equation 4.5, are calculated in Aspen Plus 10. The regressed parameters are listed in
Table 4.4. These parameters differ quite much from the parameters valid for VLE data
for the ethanol / water azeotrope, which are also listed in Table 4.4.
ln( ij )  Aij 
Bij
T
 Cij  ln(T )  Dij  T (298.14 < T < 408.65 K)
-27-
(4.5)
in which:
ij
Aij t/m Dij
T
UNIQUAC binary interaction parameter (-)
UNIQUAC regression parameters
Temperature (K)
Table 4.4: Regressed Aspen-UNIQUAC parameters of water (i) / ethanol (j) / PAA VLE data
UNIQUAC
Aji
Bij
Bji
Cij Cji Dij Dji
Aij
Aspen-default
(with azeotrope)
-2.4936
2.0046
756.95
-728.97
0
0
0
0
Regression
(without azeotrope)
-3.6339
-143.15
1,519.8
50,000
0
0
0
0
Nevertheless it can be seen from Figure 4.7 that the curve fits quite well through the
experimental data. However, because of the very small amount of experimental data
points at high ethanol mole fractions, which is near the possible azeotrope, the
regressed parameters seem not very realistic. Trying to make a feasible design of the
distillation column the regressed equilibrium curve is used.
Mole fraction ETHAN-01
0.4
0.6
0.8
1
y vs. x
0.2
Exp D-1 R-1
0
Est D-1 R-1
0.2
0.4
0.6
Mole fraction ETHAN-01
0.8
1
Figure 4.7: Aspen-UNIQUAC regression of the experimental data of ref. 2.
An additional doubt of the validity when using the regression results arises. The
regressed equilibrium curve is only valid at the pressure of 1 bar, but in the distillation
column some pressure drop has to be designed. When using the curve only at 1 bar it
reflects the influence of PAA well. But in Figure 4.8 can be seen that the influence of
pressure on the equilibrium curve is large. Here a trade-off situation occurs: should
pressure drop be designed in this case, in spite of the large influence of the
UNIQUAC parameters on pressure dependent VLE data, or should the distillation
column be designed without any pressure drop.
The choice has been made to design a pressure drop in the column because in practice
there is always some pressure drop in a column. As Figure 4.8 shows the azeotrope
does occur at higher pressures than about 1 bar. Because the purpose of this option is
to avoid the azeotrope by adding an entrainer, the pressure in the column design is
kept below 1.0 bar. In this way there will be no azeotrope and equilibria curves that
-28-
0.8
Vapor Molefrac ETHAN-01
0.4
0.6
0.2
1
0.8
Vapor Molefrac ETHAN-01
0.4
0.6
0.2
1
0.8
Vapor Molefrac ETHAN-01
0.4
0.6
0.2
1
are extrapolated will be used. Obviously there is doubt of the validity of this design,
so further experimental data should be obtained before final design and construction
are realised. To compensate slightly for these uncertainties the column is designed in
such way that the product concentration is higher than the original specification.
Y-x for WATER/ETHAN-01
1.1 bar
1.0 bar
0
0
0.9 bar
0
0.05 0.1 0.15 0.2 0.25 0.3 0.35 0.4 0.45 0.5 0.55 0.6 0.65 0.7 0.75 0.8 0.85 0.9 0.95
Liquid Molefrac ETHAN-01
Figure 4.8: The influence of pressure on the ethanol / water / PAA equilibrium curve.
-29-
1
5
Process structure & description
In this chapter the selection of unit operations, equipment, utilities and process chemicals
will be made clear on the basis of the design criteria mentioned in Paragraph 5.1. In
accordance to the choice of the equipment and the utilities the final process conditions are
defined. The processes are extendedly drawn in the process flow schemes and in this
chapter a short description is available. Finally the process yields are given. The actual
design details are treated in Chapter 8.
5.1
Design criteria
According to the assignment only one criterion has to be fulfilled, namely the ethanol
product purity has to be 99.8 vol%. This purity has to be achieved regardless of the
specified variations in the feed stream (see Paragraph 3.3.4). Another criterion, which is
tried to be fulfilled, is a recovery of at least 99 % of the pure ethanol present in the feed
stream. Besides, the costs of the total process will be kept as low as possible. To attain
these design criteria the selection of the equipment, special process condition, utilities and
process chemicals are reviewed in the subsequent paragraphs. The influences of these
choices will be shown in Paragraph 5.7.
5.2
Unit operations
5.2.1 Distillation columns
In the chosen dehydration units, a distillation column is the backbone of each design.
Distillation columns can be plate columns or packed columns. The criteria for selection
between these two possibilities are given in Table 5.1.
Table 5.1: Selection criteria and their evaluations for plate and packed columns.
Vapour-liquid contact A plate column provides a good vapour-liquid contact and is stage wise,
while the vapour-liquid contact in a packed bed column is continuous.
However the performance of a packed column is dependent on the
maintenance of good liquid and vapour distribution throughout the bed.
There is always some doubt that good distribution can be maintained
throughout a packed column.
Accuracy
Plate columns can be designed with more assurance than packed columns,
because in packed columns there is always some doubt about the good
liquid distribution. Because the high requirements of the product there is
only a small operating range. Therefore the accuracy of the design has to be
very good to ensure that the product requirements are reached.
Efficiency
A plate column provides a sufficient liquid hold-up. This provides a good
mass transfer and therefore a high efficiency. The efficiency of a plate can
be predicted with more certainty than the equivalent term for packing due to
the liquid distribution.
Pressure drop
In general the pressure drop of a packed column can be lower than of a plate
column. The pressure drop of a plate column can be held within acceptable
limits when a sufficient area and spacing is kept.
-30-
Table 5.1 (continued)
Properties of
chemicals
Economy
Ethanol and water are non-corrosive or foaming substances. Therefore is it
not necessary to choose for a packed column on the basis of a foaming or
corrosive mixture.
Plate columns with a small diameter are quite expensive because the plates
are difficult to install. Packing is much cheaper in small diameter columns.
Using the criteria above, a choice is made for plate columns because of the high accuracy
of the design. This is a very important factor, because the high requirements of purity of
99.8 vol% have to be reached. Because the liquid distribution is an important factor on the
accuracy, efficiency and vapour liquid contact, some fluctuations in the packed column
cannot be prevented. Therefore the certainty of a plate column is chosen, in spite of its
higher costs.
Not only the type of column, but also the plate contractor in the column has influence on
the overall performance. For the selection of the plate contactor the choice can be made
between three types:
- Sieve plates
- Bubble-caps
- Valve plates
Considering the costs, sieve plates have been chosen. A sieve plate is approximately 1.5
and 3 times cheaper than valves and bubble-caps respectively. Sieve plates operate
satisfactory for most applications. A disadvantage of sieves is that the operating range of
sieves is smaller than the operating ranges for bubble-caps and valves, especially at startup and shutdown conditions. Because sieves plates rely on the vapour flow through the
holes to hold the liquid on the plate, sieves can not be operated at very low vapour rates.
Bubble caps and valves have a positive liquid seal and can operate efficiently at low
vapour rates. So sieves plates are satisfactory on the condition that column weeping is
checked.
5.2.2 Condensers
The top stream of a distillation column can be condensed partially or totally, or totally
not. The type of condenser influences the heat duty and the downstream equipment.
Because the downstream equipment operates with liquid phases a total condenser has
been chosen. This implies that one part of the liquid stream is refluxed, and the other part
can be used as liquid feed for the downstream apparatus. For the condenser a fixed tube
sheet is chosen, because this is the simplest and cheapest type of exchanger.
5.2.3 Reboilers
In designing a reboiler there can be chosen between different kinds of types:
- Forced circulation
- Thermosyphon (natural circulation)
- Kettle type (no circulation)
-31-
The forced circulation reboiler is especially suitable for viscous and fouling media or is
used under high vacuum (< 0.3 bar). Non of these circumstances will occur in the process
so there has to be made a choice between the thermosyphon reboiler and the kettle
reboiler. In most cases a thermosyphon reboiler is the most economical and most used
reboiler for applications above 0.3 bar. However a disadvantage of the thermosyphon
reboiler is that the column base has to be elevated to provide hydrostatic head required for
the thermosyphon effect (ref. 33). This affects the column supporting structure and will
increase the costs. The kettle reboiler has a lower heat transfer coefficient and there is no
liquid circulation. Generally it will be more expensive than an equivalent thermosyphon
reboiler, but the costs can be decreased fairly by implementing the reboiler in the base of
the column. The costs will be in this case competitive or lower than using a thermosyphon
reboiler, while no additional column supporting structure is necessary. Therefore a kettle
reboiler is chosen in all four design options.
5.2.4 Heat exchangers
An important part of the total plant expenses comes on the account of heating and cooling
the process streams. To utilise these heat streams more efficiently, and to decrease the
overall costs, heat exchangers are used. The most commonly used type of heat-transfer
equipment is the omnipresent shell and tube exchanger. However there are more options
available as the double pipe exchanger, plate and frame exchangers, plate fin exchangers
and air coolers. A distinguish can be made between exchangers between two process
streams and a heating or cooling equipment which requires utilities.
For the process streams exchanger a choice is made for the in chemical industries
commonly used tube and shell exchangers. This is done because of the following useful
advantages (ref. 33, p. 584):
 A large surface area in a small volume
 Good mechanical layout for pressure operation
 Can be constructed from many materials with well-established fabrication techniques
 Can easily be cleaned
Due to the advantages mentioned above, the coolers of the process streams are operated
with cooling water. However more possibilities are available to cool process streams. An
interesting option is the use of air-cooled heat exchangers, because air is cheap and easily
available and there is no probability of leakage. Air-cooled heat exchangers are mostly
used in areas, where seasonal variations in ambient temperatures are relatively small.
Because a research on the climate in Lithuania falls out of the scope of the assignment, a
first choice is made to use cooling water. Moreover cooling water is available according
to the assignment and quite constant in temperature. However air coolers remain an
interesting possibility and should be investigated (see Chapter 13). For the heating of
product streams medium pressure steam of 10 bar is used. Both the coolers and heaters are
shell and tube exchangers.
All heat exchangers are operated counter-currently. This is because a counter-current heat
exchanger is more effective than a co-current heat exchanger. The fluids flowing through
the pipe are continuously changing in temperature. If the two streams are flowing in an
opposite direction, the temperature difference between the shell and tube temperature will
show less variation than in the case of co-current flow. Therefore it is possible for the
-32-
cooling liquid to leave at a higher temperature than the heating liquid, contrary to a cocurrent operated heat exchanger, where the outlet of the heating fluid must always be
higher than that of the cooling fluid. Another advantage of counter-current flow is the
extraction of a higher proportion of the heat of the hot fluid.
The factors given in Table 5.2 will determine the allocation of the fluids in the shell or in
the tubes (ref. 33). When several factors contradict the most important one will decide the
allocation.
Table 5.2: Factors determining the fluid allocation in heat exchangers.
Factor
Rule of thumb
Corrosion
The most corrosive fluid should be allocated in the tubes.
Fouling
The fluid with the highest fouling-tendency should be allocated in the tubes.
Fluid temperatures
The fluid with the highest temperature should be allocated in the tubes.
Operating pressures
The fluid with the highest pressure should be allocated in the tubes.
Flow-rates
The fluid with the highest flow-rate should be allocated in the tubes.
Process stream heat exchangers are implemented in the design to optimally use the
available heat capacity of process streams (heat integration). In this way the amount of
heating or cooling heat exchangers and their utilities are minimised.
5.2.5 Vessels
In the four designs several vessels are used: reflux accumulators and decanters. All
vessels are designed as a cylinder because this is the cheapest shape (ref. 33). The
properties of the incoming stream and the purpose of the vessel determine the position of
the vessel. For example decanters are essentially tanks, which have to give sufficient
residence time for the droplets of the dispersed phase to settle readily. For small flow
rates, which is the case in both two designs containing a decanter, a vertical cylindrical
vessel is more economical than a horizontal one. For great stream rates the decanter will
be cheaper as a horizontal vessel. Furthermore the size of the vessels is based on a chosen
(average) residence time.
5.2.6 Pumps
For the selection of the pumps distinction can be made between dynamic pumps and
positive displacement, reciprocating pumps. Positive displacement pumps are normally
used where a high Net Positive Suction Head (NPSH) is required at a low flow rate.
Because this is not the case a dynamic pump will be installed. The by far most widely
used type in chemical industry is the centrifugal pump. It is capable of pumping liquids
with very wide-ranging properties and can be constructed from a very wide range of
(corrosion resistant) materials. Therefore the centrifugal pump is perfectly capable of
handling the fluids present in the designs and is cheaper than other types of pumps.
For a final design the selection of the pump cannot be separated from the design of the
complete piping system. The complete NPSH required will be the sum of the dynamic
head due to friction losses in the piping, fittings, valves and process equipment, and any
static head due to differences in elevation. In this design the friction losses and the design
of the piping system are left out of consideration.
-33-
5.2.7 Control valves
Valves can be divided in two types: shut-off valves and control valves. In this design only
control valves are used to regulate the flow. These valves should be capable of giving
smooth control over the full range of flow. Because the valves should be automatic
controlled globe valves are used.
5.2.8 Ultrafiltration unit
In the design option extractive distillation by PAA an aqueous solution of PAA has to be
separated in a waste water stream and a PAA stream dissolved in water. The easiest and
cheapest way of doing this is using an ultrafiltration unit. This filtration membrane
permeates water (and the hydrocarbons present in low concentrations). In this way the
goal can be achieved.
5.2.9 Zeolite membrane pervaporation unit
One of the four designs contains a normal distillation that produces the 96 w% ethanol /
water azeotrope. To selectively remove the water a zeolite pervaporation membrane unit
is used. In this way the ethanol can be purified without significant loss.
5.3
Process chemicals
As mentioned in Chapter 2 an entrainer is often used to separate the azeotropic ethanol /
water mixture. In literature several possible entrainers are given. In this paragraph the
choices of the entrainers for the relevant design options are explained.
5.3.1 Azeotropic distillation by toluene
Azeotropic distillation is a traditional process, which applies as standard case in this
report. In the past the common used entrainer was benzene (ref. 22). Nowadays there are
more possible azeotropic entrainers, such as toluene, cyclohexane, diethylether and npentane. All entrainers react severely with oxidative substances possibly resulting in fire
and explosion, are very flammable and their vapour is explosive with air (ref. 7). On basis
of these properties no choice can be made. The choice is made according to componentspecific disadvantages given in Table 5.3.
-34-
Table 5.3: Pure component properties of possible azeotropic entrainers
Component
MAC-value (ppm)
Disadvantage
Toluene
40
Low MAC-value
Benzene
1
Carcinogenic
n-Pentane
600
Toxic for water in environment
Cyclohexane
250
Toxic for water in environment
Diethylether
100
Contact with hot surfaces is prohibited, especially
steam pipes
As can be seen in Table 5.3, benzene has carcinogenic properties, while n-pentane and
cyclohexane are poisonous for the environment in the event of losses. Because of the use
of steam in the reboilers it could be hazardous to use diethylether in case of leakage. In
spite of the low MAC-value toluene is less hazardous for the environment and safer in
case of leakage than the other entrainers. So toluene is chosen as the mass separating
agent in the designed azeotropic distillation.
5.3.2 Extractive distillation by gasoline
A combination of the purpose of the dehydration unit and its final application is made in
this design. Instead of producing pure ethanol, gasoline will be used to produce gasohol
directly. For this reason the entrainer for the removal of water is obviously chosen to be
gasoline. This implies that no separation of the entrainer and the desired product is
needed.
5.3.3 Extractive distillation by PAA
For the extractive distillation a very large amount of extractive entrainers is capable of
breaking the azeotrope, for example acetic acid, 2-aminoethanol, N,Ndimethylformamide, ethylene glycol and morpholine. These entrainers are already in use
and patented, but recent research has been done on polymeric entrainers (ref. 2). This
development seems to be very interesting, because of the wide availability and low costs
of the polymers. Besides, the polymeric entrainers remain in the liquid phase, so they can
be separated easily by ultrafiltration. As entrainer the polymer polyacrylic acid (PAA)
will be chosen. This entrainer has very promising potential, because it breaks the
azeotrope already at 0.45 w% polyacrylic acid added. This is remarkable because large
streams of conventional entrainers (a minimum of 30 w%) are needed.
5.4
Utilities
In the designs the following utilities, which are available at the plant site, are used:
- cooling water
- electricity
- medium pressure (MP) steam at 10 bar
- liquid nitrogen
The extended properties and costs are tabulated in the utility costs and utility summary
sheets in Appendices 6 and 7.
-35-
5.4.1 Cooling water
In all designs cooling water is used for the cooling of the process streams. Another
possibility is the use of air-cooling, but because the uncertainty of the climate in Lithuania
cooling water is preferred. The major users of cooling water are the condensers at the top
of the column.
5.4.2 Electricity
In each design electric pumps are defined. Because the size of most of the pumps is very
small, electric pumps are the cheapest. Another possibility is to execute the pump on
steam pressure. In the case of a power failure this kind of (essential) pumps are not
affected. However steam pressure pumps are more expensive and therefore in this
conceptual design only electrical pumps are used. In a finite design steam pressure pumps
should be considered as essential points for maintaining safety.
5.4.3 Medium pressure steam
For the heating in the reboilers of the distillation columns medium pressure steam is used.
Other possibilities are furnace-heating or oil-heating. Because the distillation column
bottom streams are too small both alternative possibilities are more expensive than the
one using medium pressure steam. Medium pressure steam is also used for the heating of
process streams where necessary.
5.4.4 Liquid nitrogen
This special utility is needed to condense the permeate stream of the pervaporation
membrane-unit. This stream is nearly pure water vapour at 0.008 bar and 3.8 ºC. This
stream cannot be cooled with the available cooling water of 20 ºC. Therefore liquid
nitrogen is used as cold utility in accordance with ref. 19 and ref. 36.
5.5
Final process conditions
The main objective for designing the ethanol dehydration plant, is to minimise the costs,
without losing touch with environmental and safety issues. This means that whenever
possible the conditions of the process should be at ambient pressure and at low
temperatures. In most cases these conditions are maintained, but in some cases it is not
feasible to retain this conditions. In the four design options, only the pressure in the
options extractive distillation by gasoline and extractive distillation by PAA are altered
slightly from ambient pressure.
5.5.1 Extractive distillation by gasoline
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In the case of the extractive distillation by gasoline the separation of ethanol and water in
the distillation column can be excellently achieved at ambient pressure, but then the top
temperature of the column is 30.6 ºC. This temperature complicates the condensation of
this top stream. A difference of temperature between the cooling water and the process
stream should be at least 10 ºC to maintain a heat transfer (see Chapter 8). The cooling
water comes in at a temperature of 20 ºC. To meet the requirement mentioned above, the
maximum outlet temperature of the cooling water is 20.6 ºC. This implies that an
enormous amount of cooling water is necessary to cool the process stream and the
capacity of the cooling water is not totally used. The costs to condense the top stream will
grow out of proportion. Therefore the column has to be adjusted to the available utilities.
This implies that the column has to be brought under pressure, until the top temperature is
high enough, to use the cooling water optimal. In the mean time the separation has to stay
satisfactory. This point is reached when the column is operated at a pressure of 2.3 bar.
The top stream has a temperature of 47.5 ºC, so cooling water can be used as a utility.
5.5.2 Extractive distillation by PAA
In the case of PAA there is a problem in the thermodynamics of the distillation column,
earlier discussed in Chapter 4. The regressed UNIQUAC-parameters are valid at 1 bar,
while the pressure in the distillation column changes. Furthermore it appears that the
dependence of the UNIQUAC-parameters on the pressure dependent vapour-liquid
equilibrium is strong (Figure 4.8 in Chapter 4). In the design the pressure is kept below 1
bar. This is done because the essence of the design is the avoidance of the azeotrope of
ethanol and water. The final conditions of the column are 0.84 bar at the top and 1 bar at
the bottom.
5.6
Process Flow Schemes
For each designed option a process flow scheme is added in Appendix 8. In this paragraph
each design is described according the accompanying process flow scheme. The bases for
the process flow schemes are the Aspen Plus 10 simulations. In Appendix 9 a description
of the Aspen Plus 10 files are given.
5.6.1 Azeotropic distillation by toluene
The toluene make up stream <3> (available at T = 30 °C and p = 1 bar) is added to the
ethanol feed stream <1> and is led to the heat exchanger (E01) to heat up before the
stream is pumped to distillation column (C01), which is operated at 1 bar. In the
distillation column (C01) the ternary azeotropic mixture of toluene, ethanol and water
comes overhead <8> and a polluted ethanol stream comes over the bottom <23>. The
vapour of the top stream is condensed in the total condenser (E02) and partly refluxed
(stream <10>) to the distillation column (C01). The other part (stream <11>) is led to
decanter (S01), where the toluene and the ethanol / water layer are separated. The toluene
stream <14> is fed to distillation column (C03) (1 bar) to purify the toluene from
impurities like ethyl acetate and acetaldehyde. Although the streams are very small, a
distillation column is necessary to prevent accumulation of the impurities in the process.
The top stream <16> of distillation column (C03) which contains waste water is partly
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refluxed and partly mixed with other waste water streams. The bottom stream <21> of
column (C03) contains pure toluene and is recycled to distillation column (C01). The
mass flow of toluene that stays in the process by recycling is 560 kg/h.
The bottom stream <23> of distillation column (C01) is led to distillation column (C02)
(1 bar). In this column (C02) the polluted ethanol stream <24> is purified from its
impurities isobutyl alcohol and isopentyl alcohol. The top stream <25> is partly refluxed
and partly led to heat exchanger (E01) to cool down. Further cooling of the product
stream <30>, which contains 99.9 w% ethanol, is obtained in heat exchanger (E09). The
bottom stream <32> contains waste water. This stream is mixed with the other waste
water streams <13> and <20>, respectively originated from the decanter (S01) and
distillation column (C02).
For each stream the exact composition, temperature, pressure, phase and enthalpy can be
found in the process stream summary in Appendix 10.
5.6.2 Extractive distillation by gasoline
After an increase in pressure, the ethanol feed <5> and the gasoline feed <2> (available at
T = 20 °C and p = 1 bar) are led to distillation column (C01). Before entering the
distillation column gasoline stream <2> is split into stream <3>, distillation feed, and
stream <19>, used later on to produce gasohol. In the column (C01) the ethanol is
separated from the water and the water comes overhead together with an amount of
gasoline <6>. This vapour is condensed and is partly recycled to the column (C01),
stream <8>. The other part, stream <9> is cooled down in heat exchanger (E03) before it
is led to decanter (S01). In the decanter (S01) a water stream <12> is separated from a
gasoline stream <13>. The water stream <12> is led to the water storage. The gasoline
stream <13> heated in heat exchanger (E04) before it is recycled to distillation column
(C01).
The bottom stream <16> of column (C01) contains a mixture of gasoline and ethanol,
which is cooled down in heat exchanger (E04). The mixture is mixed with gasoline stream
<20> to produce stream <21>, which is gasohol with the desired 10 w% ethanol.
For each stream the exact composition, temperature, pressure, phase and enthalpy can be
found in the process stream summary in Appendix 10.
5.6.3 Extractive distillation by PAA
The ethanol feed <1> is led to pump (P01) before it is led to heat exchangers (E01) and
(E02) to heat up. This stream <5> is mixed with the recycle stream of water and PAA
<20>. The mixed stream <6> is led to distillation column (C01), where the mixture is
separated in an ethanol stream <7> at the top and a water / PAA stream <14> at the
bottom. The top stream of the distillation column (C01) contains no polyacrylic acid. It is
condensed in heat exchanger (E03) and partly refluxed. The pressure of the ethanol stream
<10> is brought to 1 bar in pump (P03). The ethanol is decreased in temperature by heat
exchanging it with the feed stream in (E01) and in heat exchanger (E06). The product
stream <13> is led to a storage tank.
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At the bottom of the column (C01) a mixture of water and PAA <14> is increased in
pressure by pump (P04) before the stream is cooled down in heat exchanger (E02). The
stream is further cooled down in heat exchanger (E05) and led to the ultrafiltration unit
(S01). Therein 95 w% of the water and hydrocarbons is split off the PAA. The wastewater
stream <18> is led to a biological waste water treatment. In stream <20> the mass ratio
PAA / water is 39 / 61 (for transporting as recycle stream). It is decreased in pressure and
mixed with the feed stream <5>. It is assumed that the polymer is recycled completely so
in steady state no PAA will have to be added. The mass flow of PAA that stays in the
process by recycling is 5 kg/h. This is consistent with the 0.45 w% of the azeotropic
mixture mentioned in ref. 2.
For each stream the exact composition, temperature, pressure, phase and enthalpy can be
found in the process stream summary in Appendix 10.
5.6.4 Normal distillation followed by membrane purification
The ethanol feed stream <1> is heated in heat exchanger (E01) before the stream is fed to
distillation column (C01). In the column (C01) the azeotropic mixture is coming overhead
and the redundant water over the bottom. The top stream <4> is condensed and partly
refluxed. The azeotropic mixture is increase in pressure in pump (P03) and heated in heat
exchanger (E04). Further heating to 120 ºC is obtained by heater (E05) before the mixture
is led into membrane unit (S01). In the membrane unit (S01) water <11> is separated
from the ethanol by pervaporation. The ethanol stream <12> is brought back to 120 ºC by
heater (E06) and further purified in membrane unit (S02). The ethanol stream <15> is
heated once again to 120 ºC and its final purification takes place in membrane unit (S03).
The purified ethanol stream coming out from membrane unit (S03) is cooled down in heat
exchanger (E04). The ethanol stream is cooled down in heat exchanger (E09) and further
cooled down in cooler (E10) before the product stream <22> is stored.
The released water streams <14> and <17> from respectively membrane unit (S02) and
(S03) are mixed with water stream <11> from membrane unit (S01) to create stream
<26>. This vaporous, vacuum stream is condensed in heat exchanger (E08) with liquid
nitrogen. The vacuum pump (P04) is only used during start-up procedures. To increase
the pressure of stream <27> a hydrostatic pressure increase is utilised before the stream is
heated in heat exchanger (E09). This is to prevent the production of ice, see Chapter 8.
Subsequently the water stream <29> is increased to ambient pressure in pump (P05).
The water stream <23> originating from the bottom of column (C01) is cooled down in
heat exchanger (E01) and then mixed with stream <30>. The mixed stream <31> is led to
a biological waste water treatment.
For each stream the exact composition, temperature, pressure, phase and enthalpy can be
found in the process stream summary in Appendix 10.
5.7
Process performance
After the considered selections of equipment, entrainers, utilities and process conditions
in Paragraph 5.2, the purities and recoveries of the four options can be calculated. Besides
-39-
a comparison on the purity and recovery a comparison can be made between the processes
on the basis of the process yields.
5.7.1 Purity and recovery
In Table 5.4 an overview is given of the purity and recovery of the four designed
processes.
Table 5.4: Summary of purity and recovery of the designed processes
Process option
Purity (vol%)
Purity (w%)
Azeotropic distillation by toluene
Extractive distillation by gasoline
Extractive distillation by PAA
Normal distillation followed by
membrane purification
99.95
99.89
99.94
99.88
Recovery
(kg/kg)
0.955
1.000
0.999
99.86
99.84
0.990
As is shown in Table 5.4 all options attain the required minimum purity of 99.8 vol%. In
the case of extractive distillation by gasoline the required purity cannot be obtained
because the ethanol is directly mixed with the gasoline. Instead of the purity, the
maximum allowed amount of water can be compared with the actual amount of water in
the gasohol. This absolute amount is calculated by multiplying the amount of ethanol in
the feed stream with the obtained recovery of ethanol in the gasohol stream and with the
maximum percentage of impurities allowed in the ethanol content of gasohol (0.2 w%).
Calculated in this ways the maximum allowable amount of water in the gasohol product
stream becomes:
8,673 t/a ethanol in the feed stream  100 % recovery  0.2 % = 17.35 t/a = 5.5910-4 kg/s
H2O. The actual amount of water amount to 1.8.10-4 kg/s. This implies that also this
criterion is completely satisfied. Furthermore the gasohol produced contains the desired
10 w% ethanol.
The other design criterion, to obtain a recovery of 99 %, is achieved in all options except
for the azeotropic distillation by toluene. In this option only a recovery of
95.5 % is obtained. This is mainly caused by the incomplete separation in the two
distillation columns where ethanol is separated.
5.7.2 Process Yields
Process yields are important parameters for monitoring processes and compare them with
other options. Several yields of the four chosen options are compared in Table 5.5. More
detailed information can be found in Appendix 11.
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Table 5.5: Summary of process yields
Feed
Waste water
Process chemicals
Cooling water
Electricity
Steam
Azeotropic
distillation
by toluene*
Extractive
distillation
by
gasoline**
Extractive
distillation
by
gasoline*
Extractive
distillation
by PAA*
1.21
0.22
0.01
61.6
1.12
2.60
0.12
0.01
0.90
4.43
0.25
0.19
1.16
0.15
8.99
44.3
2.48
1.94
1.16
0.16
79.9
1.04
3.39
t/t product
t/t product
t/t product
t/t product
kWh/t product
t/t product
Normal
distillation
followed by
membrane
purification*
1.17
0.17
64.1
0.80
2.82
*on basis of ethanol **on basis of gasohol
To compare the values of the four options a distinction has to be made between the three
options that have ethanol as product and the extractive distillation by gasoline, which has
gasohol as final product. Therefore, for the last option an analogue is made to be able to
compare the four options.
As can be seen in Table 5.5 the azeotropic distillation by toluene is uneconomical in using
chemicals, both feed as process chemicals, to produce the ethanol. Therefore a large flow
of waste water is present. The other three options are quite comparable in the use of feed
and their amount of waste water. The yields for cooling water, electricity and steam in all
the options are in the same order of magnitude.
-41-
6
Process control
In this chapter the process control is explained for the four designs. Because of the
great similarities in the designs with respect to process control, the control systems are
described in general. The specific control systems can be viewed in the process flow
schemes in Appendix 8.
6.1
General considerations
In the designs only basic control is considered. Instruments are provided to monitor
the key process variables during plant operation. It is desirable that the process
variable to be monitored is measured directly. Overspecification of equipment or the
total process should be avoided by ensuring that two control valves are never in series
in the same pipeline.
Some process control will have effect on other equipment. This is for example the
case in a series of control valves and pumps. A valve decreases the pressure of a
liquid and causes thermodynamically vapour. If the pump is positioned behind the
control valve, the pump will suck in some vapour and cavitation in the pump will
occur. By changing the arrangement in such way that the pump is present before the
control valve, no cavitation can occur.
6.2
Control of a distillation column
To control a distillation column the number of controls may not exceed the number of
degrees of freedom to prevent overspecification of the column. The number of
degrees of freedom is the difference between the number of variables and the number
of equations. A binary distillation column has six degrees of freedom (ref. 35, p.87).
However the designed distillation columns have several components and several feed
streams and thus the influence on the number of degrees of freedom has to be
checked. It appears the number of equations and variables both increase equally by
increasing the number of components and feed streams. Therefore the number of
degrees of freedom remains the same by increasing the number of components and the
number of feed streams and thus the distillation columns still have six degrees of
freedom.
The applied controllers for the distillation column tune the following variables:
- the feed flow(s)
- the pressure in the column
- the liquid level in the reflux accumulator
- the reflux ratio
- the liquid level in the bottom of the column (reboiler)
- the temperature in the column
With these six controllers the distillation column can be controlled. The distillation
column is then exactly specified. Each controller will be discussed below.
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A flow controller controls the feed flow before it is fed to the distillation column.
When the flow is too large the controller diminishes the flow and when the flow is too
small the flow is increased, by setting the valve in the stream.
A pressure controller at the top of the column controls the pressure in the whole
column. When more vapour is condensed a larger flow is withdrawn from the column
and the pressure in the column is reduced. When less vapour is condensed the
pressure is increased. By adjusting the cooling water flow in the condenser the
pressure in the column is controlled.
A level controller in the reflux accumulator controls the liquid level in the reflux
accumulator. This liquid level controller adjusts the flow of the top product stream.
When the level rises the outflow is increased and when the level decreases the outflow
is decreased. In this way a steady liquid level can be obtained. The level controller
prevents the reflux accumulator to flood or dry up.
To control the reflux of the columns a ratio controller is installed at the top of every
column. The flow of the reflux and the flow of the top product stream are measured.
By setting the reflux ratio the flow of the reflux stream can be adjusted by a certain
flow of the top product stream. The flow of the top product stream may change due to
the level controller of the reflux accumulator.
A level controller at the base of the column controls the liquid level in the base of the
column. This liquid level controller adjusts the flow of the bottom product stream to
obtain a certain level. When the level rises, the flow is increased and when the level
decreases the flow is decreased. In this way a steady liquid level can be obtained. The
level controller prevents the base of the column to flood or dry up.
A temperature controller controls the temperature of the whole column. It controls the
temperature by adjusting the amount of steam in the reboiler. When the temperature in
the column is too low the amount of steam increases resulting in a raise of the
temperature due to the increasing amount of vapour recycled to the column. When the
temperature is too high the amount of steam is diminished to decrease the
temperature.
In Table 6.1 a summary is given of all the columns in the designs, which are
controlled with the controllers mentioned above. The position of the controllers can be
viewed in the process flow schemes, Appendix 8.
Table 6.1: Distillation columns in the four designs.
Design
Azeotropic distillation by toluene
Extractive distillation by gasoline
Extractive distillation by PAA
Normal distillation followed by membrane purification
Columns
C01, C02, C03
C01
C01
C01
In the design azeotropic distillation by toluene, the flow controllers in the recycle of
column (C01) and the feed of column (C03) are not possible. The liquid level
controllers in the reboiler of respectively column (C03) and (C01) already determine
these flows.
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6.3
Control of heaters and coolers
To make sure the heaters and coolers function as they should, a temperature controller
is placed at the outlet of the each heater respectively cooler. The controller adjusts the
amount of medium pressure steam respectively cooling water, in case it senses a
deviation in temperature.
In every cooler (including the condenser) the control valve is placed in the outlet of
the cooling water stream to prevent the emptying of the cooler when the control valve
is closed. This is done to prevent fouling. When the valve is placed in the inlet of the
cooler, the cooling water evaporates and fouls the shell or tubes of the cooler in case
of closure of the valve. Another reason is to prevent an uncontrolled cooling water
stream at low flows. In the design, normal distillation followed by membrane
purification, the control valve is placed in the inlet stream of the liquid nitrogen of
condenser (E08). This is because the inlet liquid flow is desired to be controlled and
not the outlet gas flow. Fouling is not relevant here.
The control valve in the heaters (including the reboilers) is placed in the inlet of the
medium pressure steam. In practice it makes no difference if the control valve is
placed in the inlet or in the outlet of the stream of medium pressure steam. So the
choice to control the inlet stream is completely arbitrary. In Table 6.2 the heaters and
coolers are summarised for each design. The positions of the controllers can be
viewed in the process flow schemes, Appendix 8.
Table 6.2: Heaters and coolers.
Design
Azeotropic distillation by toluene
Extractive distillation by gasoline
Extractive distillation by PAA
Normal distillation followed by
membrane purification
6.4
Heaters
E03, E05, E07
E02
E04
Coolers
E02, E04, E06, E08, E09
E01, E03
E03, E05, E06
E03, E05, E06, E07
E02, E08, E10
Control of a decanter
To control the separation in the decanters a level controller is placed. The level
controller measures and controls the height of the liquid-liquid-layer located between
the two phases in the decanter. When the water phase level has to be decreased or
increased, the water flow is adjusted to reach the required level. In the azeotropic
distillation by toluene and in the extractive distillation by gasoline a decanter is used.
In Table 6.3 the decanters with level controllers are summarised. The position of the
controllers can be viewed in the process flow schemes, Appendix 8.
Table 6.3: Decanters.
Design
Azeotropic distillation by toluene
Extractive distillation by gasoline
Decanter
S01
S01
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6.5
Control of the ultrafiltration unit in the extractive distillation by PAA
The retentate flow of the ultrafiltration unit (S01), in the extractive distillation by
PAA, is measured and controlled. The retentate flow over the ultrafiltration unit is
needed to control the flux through the filter. When the flux trough the filter decreases
the retentate flow will increase. This is an unwanted effect. So to restore the flux the
control valve decreases the retentate flow. When the retentate flow is decreased the
pressure over the filter is increased resulting in an increased flux.
6.6
Control of the membrane unit in the normal distillation followed by
membrane purification
A flow controller controls the flow of the permeate flow of the last membrane module
(S03). The flow controller is needed to assure a flux over the membranes. When the
permeate flow decreases the retentate flow will be decreased by closing the valve to
increase the pressure difference and so the flux. When the permeate flow increases the
retentate flow is increased to decrease the pressure difference and the flux again.
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7
Mass and heat balances
In this chapter the mass and heat balances of the four designs are checked. First the
mass and heat balances of the azeotropic distillation by toluene are summarized.
These are followed by those of the extractive distillations by gasoline and polyacrylic
acid. Finally the mass and heat balances of the normal distillation followed by
membrane purification are represented.
7.1
Mass and heat balances of the azeotropic distillation by toluene
In Table 7.1 the mass and heat balances for the azeotropic distillation by toluene are
presented. As can be seen the design is in mass balance. There is no mass production
or consumption. The enthalpy of the process however is not in heat balance. This is
because the values of the enthalpy of the inlet and outlet streams and the duties of the
equipment are rounded off in this table. In Appendix 12 the mass and heat balances
are represented for the individual equipment. The values used in Appendix 12 show
that the process is in mass and heat balance. In Appendix 12 the values have more
significant digits than represented. The totals in the spreadsheet are made by taking
the sum of the not rounded numbers and are more accurate than the values in Table
7.1.
Table 7.1: Mass and heat balance of the azeotropic distillation by toluene
IN
OUT
Mass (kg/s)
Enthalpy (kW)
Mass (kg/s)
Enthalpy (kW)
Inlet stream(s)
0.327
-2,356
Outlet stream(s)
0.327
-2,351
Equipment total
1,592
1,588
Total
0.327
-763
0.327
-763
* The sum of the enthalpy IN is not the same as the displayed total value due to the rounding off of the
displayed values.
In Table 7.2 the mass streams per component are shown for the in- and outlet streams
of the battery limit. The third column in Table 7.2 is empty because all components
are in mass balance (zeros are not shown). In Appendix 10 the process stream
summary is given. The final table in this Appendix shows a component and total mass
and total heat balance over the plant.
Table 7.2: Overall component mass balance.
IN
OUT
OUT-IN
Component
Mass (kg/s)
Mass (kg/s)
Mass (kg/s)
Ethanol
0.280
0.280
Water
0.042
0.042
Toluene
0.003
0.003
Isopentyl alcohol
0.002
0.002
Isobutyl alcohol
0.001
0.001
Ethyl acetate
0.000
0.000
Acetaldehyde
0.000
0.000
Total*
0.327
0.327
* The sum of the component mass streams is not the same as the total mass stream given in Table 7.1
and Table 7.2. This is due to the rounding off of the displayed values.
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7.2
Mass and heat balances of the extractive distillation by gasoline
In Table 7.3 the mass and heat balances for the extractive distillation by gasoline are
presented. As can be seen the design is in mass and heat balance. No mass or heat is
produced or consumed. In Appendix 12 the mass and heat balances are represented
for the individual pieces of equipment.
Table 7.3: Mass and heat balance of the extractive distillation by gasoline
IN
OUT
Mass (kg/s)
Enthalpy (kW)
Mass (kg/s)
Enthalpy (kW)
Inlet stream(s)
2.837
-7,030
Outlet stream(s)
2.837
-6,978
Equipment total
1,266
1,215
Total
2.837
-5,763
2.837
-5,763
* The sum of the enthalpy IN is not the same as the displayed total value due to the rounding off of the
displayed values.
In Table 7.4 the mass streams per component are shown for the in- and outlet streams
of the battery limit. The third column in Table 7.4 is empty because all components
are in mass balance (zeros not shown). In Appendix 10 the process stream summary is
given. The final table in this Appendix shows a component and total mass and total
heat balance over the plant.
Table 7.4: Overall component mass balance.
IN
Component
Mass (kg/s)
Ethanol
0.280
Water
0.042
Toluene
0.251
Ethyl acetate
0.000
Isobutyl alcohol
0.001
Isopentyl alcohol
0.002
Acetaldehyde
0.000
1-Hexene
0.151
2-Methyl-2-butene
0.101
Methylcyclopentane
0.327
Methylcyclohexane
0.176
N-pentane
0.377
N-hexane
0.302
2-Methylpentane
0.302
3-Methylpentane
0.151
N-Heptane
0.176
2-Methylbutane
0.201
Total*
2.837
OUT
Mass (kg/s)
0.280
0.042
0.251
0.000
0.001
0.002
0.000
0.151
0.101
0.327
0.176
0.377
0.302
0.302
0.151
0.176
0.201
2.837
OUT-IN
Mass (kg/s)
* The sum of the component mass streams is not the same as the total mass stream given in Table 7.3
and Table 7.4. This is due to the rounding off of the displayed values.
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7.3
Mass and heat balances of the extractive distillation by PAA
In Table 7.5 the mass and heat balances for the extractive distillation by PAA are
presented. As can be seen the design is in mass balance. There is no mass and no heat
production. In Appendix 12 the mass and heat balances are represented for the
individual pieces of equipment.
Table 7.5: Mass and heat balance of the extractive distillation by PAA
IN
OUT
Mass (kg/s)
Enthalpy (kW)
Mass (kg/s)
Enthalpy (kW)
Inlet stream(s)
0.324
-2,356
Outlet stream(s)
0.324
-2,352
Equipment total
2,162
2,158
Total
0.324
-194
0.324
-194
In Table 7.6 the mass streams per component are shown for the in- and outlet streams
of the battery limit. The third column in Table 7.6 is empty because all components
are in mass balance (zeros are not shown). In Appendix 10 the process stream
summary is given. The final table in this appendix shows a component and total mass
and total heat balance over the plant.
Table 7.6: Overall component mass balance.
IN
Component
Mass (kg/s)
Ethanol
0.280
Ethyl acetate
0.000
Acetaldehyde
0.000
Water
0.042
Isobutyl alcohol
0.001
Isopentyl alcohol
0.002
PAA
Total
0.324
OUT
Mass (kg/s)
0.280
0.000
0.000
0.042
0.001
0.002
OUT-IN
Mass (kg/s)
0.324
* The sum of the component mass streams is not the same as the total mass stream given in Table 7.5
and Table 7.6. This is due to the rounding off of the displayed values.
7.4
Mass and heat balances of the normal distillation followed by membrane
purification
In Table 7.7 the mass and heat balances for the normal distillation followed by
membrane purification are shown. As can be seen the design is in mass balance. No
heat or mass is produced or consumed. In Appendix 12 the mass and heat balances are
represented for the individual pieces of equipment.
-48-
Table 7.7: Mass and heat balance of the normal distillation followed by membrane
purification.
IN
OUT
Mass (kg/s)
Enthalpy (kW)
Mass (kg/s)
Enthalpy (kW)
Inlet stream(s)
0.324
-2,356
Outlet stream(s)
0.324
-2,352
Equipment total
1,772
1,768
Total
0.324
-584
0.324
-584
In Table 7.8 the mass stream per component are shown for the in- and outlet streams
of the battery limit. The third column in Table 7.8 is empty because all components
are in mass balance (zeros are not shown). In Appendix 10 the process stream
summary is given. The final table in this appendix shows a component and total mass
and total heat balance over the plant.
Table 7.8: Overall component mass balance.
IN
Component
Mass (kg/s)
Ethanol
0.280
Water
0.042
Isopentyl alcohol
0.002
Isobutyl alcohol
0.001
Ethyl acetate
0.000
Acetaldehyde
0.000
Total*
0.324
OUT
Mass (kg/s)
0.280
0.042
0.002
0.001
0.000
0.000
0.324
OUT-IN
Mass (kg/s)
* The sum of the component mass streams is not the same as the total mass stream given in Table 7.7
and Table 7.8. This is due to the rounding off of the displayed values.
-49-
8
Process and equipment design
In this chapter decision points for the process and equipment are explained. The
equipment is sized to reach the design criteria.
8.1
Integration by process simulation
As already mentioned, the process simulations are done with the flowsheet calculation
program Aspen Plus 10. It is a tool used by engineers to model any type of process for
which there is a continuous flow of materials and energy from one processing unit to
the next unit. Aspen Plus 10 can be used to model processes in chemical and
petrochemical industries, such as petroleum refining, synthetic fuels, power
generations and metals and minerals. Flowsheet models are used at all stages in the
life cycle of a process plant. Input to the model consists of information normally
contained in the process flowsheet. Output from the model is a complete
representation of the performance of the plant (ref. 4).
Because the designs of the options consist of relatively simple, familiar separation
units all modelling will be done in Aspen Plus 10. In each design a distillation column
is the backbone of the process. For the modelling of the distillation column initially
the short-cut model DSTWU is simulated. This model performs a Winn-UnderwoodGilliand shortcut design for a single feed, two-product distillation column with a total
condenser. These calculations give an estimate for the number of stages, the reflux
ratio and the feed stage. For the rigorous design of the distillation column a
RADFRAC-model is used. This is a rigorous model for all types of fractionation and
is assuming equilibrium at all stages.
As explained in Chapter 4 the thermodynamic model used is UNIQUAC. Also stated
there is the use of adapted UNIQUAC parameters in the extractive distillation by
PAA. These parameters are only applied to the distillation column (C01). The use of
these adapted UNIQUAC parameters makes it possible to run the Aspen model
without virtually adding polyacrylic acid as a component. Every other equipment
handling a stream containing PAA is simulated with the default Aspen-UNIQUAC
parameters.
In the Aspen model of the extractive distillation by gasoline the entrainer gasoline is
modelled with the eleven components in practice most occurring in gasoline. The
composition of the gasoline feed stream as modelled in Aspen Plus 10 is tabulated in
Chapter 3, Table 3.4.
For designing the heat exchangers distinguish is made between the heaters and coolers
and heat exchangers between two process streams. For the heaters and coolers with
respectively steam and cooling water the HEATER model in Aspen Plus 10 is
especially suitable. For a process stream heat exchanger a HEATX-model is used.
This is a two-stream heat exchanger, which can model counter-current streams. The
outlet conditions of one of the streams can be specified.
The ultrafiltration unit in the option extractive distillation by PAA is calculated by the
program MathCad Plus 6.0 and on basis of the obtained data simulated in Aspen Plus
-50-
10 by a separator. Because Aspen Plus 10 cannot simulate filtration units or
membranes the ultrafiltration unit is calculated in MathCad Plus 6.0. This is a blank
worksheet on which equations, graph data, text or functions, can be entered and
annotated with text anywhere on the page.
The zeolite NaA pervaporation membrane unit of the option normal distillation
followed by membrane purification is modelled in Aspen Plus 10 as a series of heaters
and separators. The unit where the liquid feed stream loses temperature because of the
pervaporation of water is simulated as follows: First the feed stream is splitted into a
permeate and a retentate stream. Then the retentate stream is reduced in temperature
because of the heat needed for pervaporation. This heat needed to permeate water and
some promille ethanol is calculated using a constant heat of evaporation of water. The
amount of ethanol is neglected for that purpose. Finally the permeate is brought to the
desired conditions. This is illustrated in Appendix 9. The compositions of the retentate
and permeate streams of each of the three membranes are designed in MathCad Plus
6.0, like the ultrafiltration unit mentioned above.
8.2
Equipment selection and design
In this paragraph the selection and calculation of the process equipment is explicated.
In Appendix 13 calculation examples are provided for the summarised values in this
paragraph.
8.2.1 Distillation columns
As mentioned in Chapter 5 each distillation column is a sieve column. These columns
can only be operated within certain limits, which are formed by irregularities such as
excessive entrainment, flooding, downcomer back-up limitation, weeping and coning.
These irregularities influence the performance and the efficiency of the column. In a
small diameter column (less than 1,2 m) each plate is fabricated complete with
downcomer and joined to the plate above and below in stacks of approximate 10
sieves. These plates are installed as stacks in the column, because the diameter is to
small for a manhole.
After choosing a plate spacing and calculating the accompanying column diameter the
sieve plate can be designed. In Table 8.1 the number of stages and the tray spacing
(HETP) is summarised for each column.
-51-
Table 8.1: Geometry of the column and the plates.
Azeotropic Azeotropic Azeotropic
distillation distillation distillation
by toluene, by toluene, by toluene,
C01
C02
C03
Number of
stages
HETP (m)
L (m)
D (m)
Hole area
Hole size
(mm)
Weir
height
(mm)
Downcomer area
Extractive
distillation
by
gasoline,
C01
Extractive
distillation
by PAA,
C01
Normal
distillation
followed by
membrane
purification,
C01
100
0.3
35.2
1.056
10%
18
0.5
14.0
0.508
10%
11
0.5
10.5
0.456
10%
10
0.6
10.9
1.052
10%
33
0.6
24.7
1.203
10%
23
0.6
18.7
0.986
10%
3
5
3
2.5
5
5
40
40
40
40
40
40
12%
12%
12%
12%
12%
12%
The length of the column is calculated with the number of stages and the tray spacing.
The base of the column is 3.0 m and the space at the top is 2.5 m. In the base of the
column the kettle reboiler is present. Because the column has plates a pressure drop in
the column is present. The column is designed in such a way that the irregularities
such as excessive entrainment, flooding, downcomer back-up limitation and coning
are not present. This is verified in Aspen Plus 10. In the columns weeping does not
take place, this is verified in Appendix 14. The specifications of each column are
presented in Appendices 15 till 18.
8.2.2 Condensers
In all condensers at the top of a distillation column cooling water is used as utility as
explained in Chapter 5. This cooling water is available at 20 ºC and is discharged at
maximally 40 ºC. With the in- and outgoing temperatures of the cooling water and
process streams the amount of cooling water in each condenser of the column is
calculated. These amounts are summarised in Table 8.2. The overall heat transfer
coefficients, U, for the condensers are determined with Figure 12.1 in ref. 33 (p. 582)
With these overall heat transfer coefficients the condenser areas can be calculated. To
make a realistic design of the condensers of the columns, a heat loss of 5 % on the
duty calculated by Aspen Plus 10 is posed. The calculations of the amounts of cooling
water and the exchange areas include this heat loss. In Table 8.2 the overall heat
transfer coefficients and the areas of the condensers are given. In Appendix 14 the
areas of the condensers are calculated. In the equipment specification sheets,
Appendices 15 till 18, the specifications of the condensers are summarised. Condenser
(E08) in the normal distillation followed by membrane purification is not a shell and
tube condenser at the top of the column. It has a spiral shape and condenses the
membrane permeate stream with liquid nitrogen. This will be discussed in Paragraph
8.2.9.
-52-
Table 8.2: Mass flow cooling water, heat transfer coefficients and areas of the condensers.
U (W/(m2.K))
Design
Condenser
A (m2)
m, cooling water
(kg/s)
Azeotropic distillation by toluene E02 of C01
11.090
600
51.6
E04 of C02
3.617
600
13.0
E06 of C03
1.471
600
5.73
Extractive distillation by gasoline E01 of C01
11.793
550
181.9
Extractive distillation by PAA
E03 of C01
22.129
600
88.0
Normal distillation followed by
E02 of C01
17.280
600
62.3
membrane purification
E08
0.1*
1000
0.62
* Nitrogen instead of cooling water
The cylindrical shell and tube condensers are placed horizontally to prevent flooding.
The vapour velocities in the designs are large in proportion to the liquid velocity.
When vertical tubes are used, flooding occurs (see Appendix 14).
In Table 8.3 the dimensions of the condenser tubes of the columns are summarised.
According to standard dimensions the wall thickness is chosen to be 2 mm (ref. 33, p.
588). With these dimensions the number of tubes needed is determined. The
condenser medium has only one pass per tube.
Table 8.3: Dimensions of condenser tubes.
dinner (m)
douter (m)
Ltube (m)
0.021
0.025
3.66
The tube pitch is 1.25 times the outside diameter. A square pattern is used for easy
cleaning of the tubes (ref. 33, p. 589). Subsequently the bundle diameter can be
calculated (ref. 33, p. 591). The shell diameter is set to be 1.1 times the bundle
diameter (ref. 33, p. 590). In Table 8.4 the values of the bundle and shell diameter are
summarised.
Table 8.4: Dimensions of the condensers.
Design
Condenser
Azeotropic distillation by toluene
E02 of C01
E04 of C02
E06 of C03
Extractive distillation by gasoline
E01 of C01
Extractive distillation by PAA
E03 of C01
Normal distillation followed by
membrane purification
E02 of C01
Dbundle (mm)
528
284
195
933
672
Dshell (mm)
580
313
214
1026
739
574
632
In Appendix 7 the utilities of all the equipment per design are summarised. In this
appendix the duties of the equipment are the duties that are really transferred to the
process flows. The amounts of medium pressure steam and cooling water are
calculated with a heat loss of 5 %. So the calculated amount of utility is based on a
larger heat transfer.
-53-
8.2.3 Reboilers
The utility used in the reboilers is medium pressure steam of 10 bar. The amount of
steam needed to reboil the liquid in the base of the column, is based on the assumption
that the medium pressure steam is condensed first and then cooled till the temperature
is 10 ºC above the boiling temperature of the mixture in the column. This is to prevent
the use of large amounts of steam. The needed amounts of steam are listed in Table
8.5. The temperature difference of 10 ºC is needed to maintain the driving force. The
overall heat transfer coefficient, U, of the reboilers is estimated at 1000 W/(m2K).
The overall heat transfer coefficient is also calculated with estimated values for the
film coefficients and fouling factors. The estimated value of 1000 W/(m2K) was close
enough to the calculated values. With the overall heat transfer coefficient the areas of
the reboilers can be calculated (see Appendix 14). In the Table 8.5 the areas are
summarised. Also in the design of the reboilers a heat loss of 5 % on the duty
calculated by Aspen Plus 10 is posed for the calculations of the amounts of medium
pressure steam and the exchange areas. In the equipment specification sheets,
Appendices 15 till 18, the specifications of the condensers are summarised.
Table 8.5: Mass flow steam, heat transfer coefficients and areas of the reboilers.
Design
Reboiler
U (W/(m2.K))
m, steam(kg/s)
Azeotropic distillation by toluene E03 of C01
0.476
1000
E05 of C02
0.148
1000
E07 of C03
0.071
1000
Extractive distillation by gasoline E02 of C01
0.542
1000
Extractive distillation by PAA
E04 of C01
0.946
1000
Normal distillation followed by
membrane purification
E03 of C01
0.757
1000
A (m2)
15.8
4.90
2.82
17.3
32.2
26.0
The reboiler is a kettle reboiler with U-tubes. In Table 8.6 the characteristics of the Utubes are summarised. According to standard dimensions the wall thickness is chosen
to be 2.6 mm (ref. 33, p. 588). With these dimensions the number of U-tubes needed
is determined. Because U-tubes are used, the medium has two passes per tube.
Table 8.6: Dimensions of reboiler U-tubes
dinner (m)
douter (m)
LU-tube (m)
0.025
0.030
4.80
The bundle diameter, Db, is determined by a sketch, see Appendix 19 for an example.
The tubes in the bundle are arranged in a square pitch, with the tube pitch of 1.5douter.
The shell diameter, Ds, is taken twice as large as the bundle diameter (ref. 33, p. 690).
In Table 8.7 the bundle and shell diameters of the reboilers are summarised.
-54-
Table 8.7: Dimensions of the reboilers.
Design
Reboiler
Azeotropic distillation by toluene E03 of C01
E05 of C02
E07 of C03
Extractive distillation by gasoline E02 of C01
Extractive distillation by PAA
E04 of C01
Normal distillation followed by
membrane purification
E03 of C01
Dbundle (mm)
510
310
240
600
780
Dshell (mm)
1020
620
480
1200
1560
690
1380
In Appendix 7 the utilities of all the equipment per design are summarised. In this
appendix the duties of the equipment are the duties that are really transferred to the
process flows. The amounts of medium pressure steam and cooling water are
calculated with a heat loss of 5 %. So the calculated amount of utility is based on a
larger heat transfer.
8.2.4 Heat exchangers
The heat exchangers are modelled by Aspen Plus 10 and this program can not include
a heat loss of 5 % in the duty transferred or the exchange area. In Appendix 15 till 18
the heat exchangers are specified. In Table 8.8 the overall heat transfer coefficient and
the heat exchanger areas are summarised.
Table 8.8: Overall heat transfer coefficients and areas of the heat exchangers
Design
Heat exchanger
U (W/(m2K))
Azeotropic distillation by toluene
E01
300
Extractive distillation by gasoline
E04
300
Extractive distillation by PAA
E01
300
E02
500
Normal distillation followed by
E01
500
membrane purification
E04
300
E09
500
A (m2)
7.0
15.4
6.88
0.58
0.54
6.01
0.082
According to the criteria mentioned in Table 5.2, Chapter 5, the fluids are allocated to
the shell and tubes. In Table 8.9 the fluids in the shell and tubes are summarised.
Table 8.9: Fluids in the shell and tubes of the heat exchangers.
Design
Heat exchanger
Tube-side
Azeotropic distillation by toluene
E01
Toluene / EtOH
Extractive distillation by gasoline
E04
Gasoline / EtOH
Extractive distillation by PAA
Normal distillation followed by
membrane purification
E01
E02
E01
E04
E09
-55-
EtOH / H2O
Water
EtOH / H2O
EtOH
EtOH
Shell-side
EtOH
2-M-butane /
EtOH
EtOH
EtOH / H2O
Waste water
H2O / EtOH
Water
8.2.5 Heaters and coolers
To make a realistic design of the heaters and the coolers, a heat loss of 5 % on the
duty calculated by Aspen Plus 10 is posed for the calculations of the amounts of
medium pressure steam, cooling water and the exchange areas. In Table 8.10 the
amounts of medium pressure steam and cooling water are summarised, in Appendix
14 the amounts of these utilities and the exchange areas are calculated. In Appendix
13 a calculation example is given. The minimum driving force for exchanging heat,
the temperature difference between the two fluids or gas and fluid, is set at 10 ºC. This
is an engineering rule of thumb (ref. 33), so some discrepancy to this rule is tolerated.
In Appendices 15 till 18 the heaters and coolers are specified. In Table 8.10 the
overall heat transfer coefficient and the exchange areas are summarised.
Table 8.10: Mass flows of the utilities, heat transfer coefficients and areas in the heaters and
coolers.
Design
Cooler Heater
U
A
m, cooling m, steam
2
2
)
(m
(kg/s) (W/(m K))
water (kg/s)
Azeotropic distillation
E08
0.070
850
0.4
by toluene
E09
0.156
500
1.6
Extractive distillation
by gasoline
E03
0.590
500
6.5
Extractive distillation
E05
0.048
1000
0.22
by PAA
E06
0.163
500
1.66
Normal distillation
E05
0.013
1000
0.37
followed by membrane
E06
0.006
1000
0.19
purification
E07
0.006
1000
0.19
E10
0.483
500
4.7
In Appendix 7 the utilities of all the equipment per design is summarised. In this
appendix the duties of the equipment are the duties that are really transferred to the
process flows. The amounts of medium pressure steam and cooling water are
calculated with a heat loss of 5 %. So the calculated amount of utility is based on a
larger heat transfer.
8.2.6 Vessels
In Table 8.11 all used vessels are listed. The vessels are shaped cylindrically. The
dimensions of the vessels are determined by setting the length to diameter proportion
and the space-time of the vessel. The length to diameter proportion is set to two and
the space-time to five minutes (ref. 33, page 401). A safety margin of 1.5 in volume of
the vessels of the reflux accumulators is posed. This safety margin is not posed on the
volume of the decanters because the decanters are completely filled with liquids. In
the reflux accumulators the space above the liquid is filled with vapour. In Appendix
13 a calculation example is made.
-56-
Table 8.11: Vessels with their space time and dimensions
Design
Vessels
Space time (s)
Azeotropic distillation by toluene
V01
5
V02
5
V03
5
S01
5
Extractive distillation by gasoline
V01
5
S01
5
Extractive distillation by PAA
V01
5
Normal distillation followed by
membrane purification
V01
5
H (m)
1.54
1.10
0.86
0.72
2.08
1.4
2.01
D (m)
0.77
0.55
0.43
0.36
1.04
0.72
1.00
1.80
0.90
8.2.7 Pumps
When a process stream needs to be pumped up to a higher pressure, the capacity of
the pump is determined by the duty calculated by Aspen Plus 10. Because the feed
streams are taken from a storage tank the pumps also need a certain amount of energy
to get over the difference in height between the storage tanks and the inlet of the
columns. The pumps in the recycle streams and in the reflux streams also need this
capacity to get over the difference in height.
To determine the difference in height it is assumed that the piping is 0.5 m above the
ground level. The capacity of the pumps is then the sum of the duty determined by
Aspen Plus 10 and the duty for the difference in height of the above calculation. A
pump efficiency of 50 % is taken into account. In Appendix 13 a calculation example
is made.
In Table 8.12 the pumps and their capacity are summarised. In Appendices 15 till 18
the pumps are specified.
Table 8.12: Pumps with their capacities in the designs.
Design
Feed pumps
No.
Azeotropic distillation by toluene
Extractive distillation by gasoline
Extractive distillation by PAA
P01
P02
P05
P07
P01
P02
P03
P01
Capacity
(kW)
0.021
0.183
0.024
0.085
0.569
0.638
1.096
0.058
Reflux pumps
No.
P03
P04
P06
Capacity
(kW)
0.746
0.032
0.001
P04
0.189
P02
0.952
Pressure increase
pumps
No.
Capacity
(kW)
P03
0.013
P04
0.020
Normal distillation followed by
P01
0.060
P02
0.470
P03
0.267
membrane purification*
P05
0.002
*
Pump (P04) is not listed here because this vacuum pump is not needed in normal operation.
-57-
8.2.8 Ultrafiltration unit
In Chapter 5 it is already mentioned that an ultrafiltration membrane is used to
separate the polymer PAA from the water. The ultrafiltration membrane is designed
according to ref. 31. In Table 8.13 the designed properties of the membrane are listed.
Table 8.13: Ultrafiltration unit properties.
Property
Material
Pore size
Inlet pressure
Outlet pressures retentate / permeate
Temperature
Permeate flux
Mass-split fraction permeate / retentate PAA
Mass-split fraction permeate / retentate water
and hydrocarbons
Membrane area
Value
A closer to be specified polymer,
for example polypropylene
10-8 m
3 bar
3 bar / 1 bar
40 ºC, isothermic
30 kg/(m2.h)
0
0.95
5.31 m2
The membrane area is calculated with the available permeate flux, which is an
intrinsic property of the membrane unit, and the desired split fraction. The goal is that
95 w% of the water and hydrocarbons permeate, while all PAA will not permeate
because the molecules are too large to permeate through the membrane pores. The
result is that the permeate stream, in which PAA of molecular weight 2,000 kg/kmole
is most concentrated, has a mass-ratio of 39 PAA over 61 water. It is assumed that
PAA is completely dissolved in water. This assumption is made on the basis of
literature (ref. 3) which says: The maximum solubility mass-ratio of PAA with
molecular weight 100,000 kg/kmole is 35 PAA over 65 water. Because used PAA has
a much lower molecular weight than the one in literature it is assumed PAA will be
dissolved and no significant ‘cake-formation’ on the membrane surface will take
place.
8.2.9 Zeolite NaA pervaporation membrane unit
To purify the ethanol / water azeotrope without loss of ethanol a pervaporation
membrane unit is used in the design option normal distillation followed by membrane
purification. Water and some promilles ethanol permeate through the membrane. It is
assumed that ethyl acetate and acetaldehyde do not permeate, because both substances
are hydrophobic and the membrane is of a hydrophilic character. Both isobutyl and
isopentyl alcohol, which occur in actually negligible amounts, are also not assumed to
permeate.
One membrane unit consists of 3 identical tube membrane modules. The properties of
one module (ref. 19) are given in Table 8.14. The pressure drop in the tubes is
negligible. Calculations of the pressure drop are given in Appendix 13.
-58-
Table 8.14: Properties of one tube membrane module.
Property
Value
OD
Tube dimensions
12 mm x 9 mmID x 800 mmL
2
Membrane area per tube (m )
0.03
Number of tubes
148
2
Total membrane area
(m )
4.44
Membrane material
Zeolite NaA on a 65 w% over 35 w%
alumina / silica support
The membrane area is calculated with the given stream magnitude and the designed
permeate flux magnitude. The pervaporation membrane properties are listed in Table
8.15.
Table 8.15: Properties of pervaporation membrane unit.
Property
Separation factor
Permeate flux
(kg/(m2.h))
Total membrane area
(m2)
Value
5600
4.3
13.2
The permeate flux is valid at an (almost) saturated liquid feed of 120 ºC. Because of
temperature loss due to pervaporation, the retentate stream must be heated to 120 ºC
before entering the next membrane.
Special attention is paid to the permeate side. A constant heat of vaporisation of
2202.2 kJ/kg at 120 ºC is used for the pervaporation of water (ref. 34, p.670). The
water vapour permeate is at the conditions of 0.008 bar and 3.8 ºC (saturated vapour).
The vacuum pressure is maintained by condensing the water vapour with nitrogen.
Cooling water is too hot to use as utility. By condensing the vapour the volume of the
stream decreases enormously and therefore the pressure decreases enormously too.
This suction by condensation is shown in Figure 8.1.
Figure 8.1: Vacuum suction by condensing the water vapour with nitrogen.
A vacuum pump is installed as back-up when permeate pressure is too high, but in
normal operation this pump is not necessary. Also it must be used in a start-up of the
process after some idle time. After condensing the water pressure is built up in a
vertical pipeline. This is necessary to prevent the making of ice. In Figure 8.2 the
phase diagram of water is shown. This makes clear no ice will be made at stated
conditions.
-59-
Figure 8.2: Phase diagram of water.
8.3
Special issues
8.3.1 Construction material
For the selection of the construction material not only the mechanical properties of the
material, like strength and toughness are taken into account, but also other aspects like
the properties of the process fluid, corrosion and costs are important. Especially
corrosion can cause special requirements to the construction materials. Corrosion can
occur in a variety of ways, like pitting, galvanic, intergranular and crevice corrosion.
In this paragraph only the uniform corrosion due to general wastage of material will
be taken into account. For a thoroughly investigation of corrosion and the resistant
materials an expert should be informed.
Because the process temperatures and pressures are moderate and the ethanol / water
is a non-corroding system, carbon steel can be used as construction material. (ref. 28,
p.23-26 t/m 23-27) Carbon steel is the most common, cheapest and most versatile
metal used in industry. It has excellent ductility, permitting cold operations and is
very weldable. Carbon steel is easily the most commonly used material in process
plants despite its somewhat limited corrosion resistance. It is routinely used for most
organic chemicals and neutral or basic aqueous solutions at moderate temperatures.
Heated for prolonged periods at temperatures above 455 ºC the carbon may be subject
to segregation of carbon to graphite. This undermines the strength of carbon steel. The
corrosion rate will be dependent on the temperature and concentration of the corrosive
fluid. The expected corrosion rate is approximate 0.25 mm/year. This is according to
ref. 33 (p. 251) completely satisfactory.
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Carbon steel will be used for the columns (except for the trays), decanters, coolers,
heaters, heat exchangers and pumps (except for the rotors). For the rotors in the
pumps and the sieves in the columns stainless steel is used. The high stream velocity
causes the need for greater strength. Generally stainless steel is iron-based, with 12 to
30 % chromium and 0 to 22 % nickel. The material is heat- and corrosion-resistant,
non-contaminating and easily fabricated into complex shapes. The expected corrosion
rate of stainless steel is approximate 0.5 mm/year.
8.4
Equipment data sheets
In Appendices 15 till 18 the equipment data sheets for all equipment in the designs are
presented. The appendix begins with the equipment data summary sheets. These
sheets provide quick reference for overall dimensions and process conditions. After
these summaries the equipment data specification sheets are presented. These sheets
provide the design details.
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9
Wastes
In this chapter the direct wastes of the four design alternatives are described. First it is
made clear what these wastes and their toxicological properties are. Secondly the
influence of upstream processes and process units on the wastes are explained.
9.1
Identification of wastes
The wastes of each design option can be divided in liquid and solid wastes. The liquid
wastes are:
- Process waste water; most of the water withdrawn from the ethanol feed stream
- Utility water:

cooling water; heated up in a condenser or cooler

condensed steam in reboiler or heater
The solid wastes are:
- PAA; in steady state operation all PAA stays in the process so no PAA has to be
added. It is taken into account that every two years some PAA will be taken from
the ultrafiltration unit and new PAA will be added.
- Membranes; when exhausted after 3 years (pervaporation- and ultrafiltration-unit)
In this conceptual design it is further assumed that cooling water can be disposed of at
40 ºC. From this point of view cooling water is not considered as waste. The solid
wastes PAA and membranes are occasional and in this conceptual design not further
dealt with. This is because costs of this waste disposal is unknown and assumed to be
not of significant influence on the economy.
The main waste is the continuous waste water stream coming out of each designed
plant option. These streams contain most of the water of the feed stream and some
quantities of the other components present in the feed stream. Table 9.1 shows the
composition of the waste water stream of every design option. These values are taken
from Appendix 10.
Table 9.1: Composition of waste water stream of every design alternative.
Component
Azeotropic
Extractive
Extractive Normal distillation
distillation by
distillation by
distillation
followed by
toluene (kg/h) gasoline (kg/h)
by PAA
membrane
(kg/h)
purification (kg/h)
Ethanol
45.07
0.96
10.38
Water
150.37
150.31
150.95
150.54
Isopentyl alcohol
5.44
5.44
5.44
Isobutyl alcohol
1.79
1.81
1.81
Ethyl acetate
0.73
0.00
0.00
Acetaldehyde
0.44
0.00
0.00
Toluene
12.10
Total
215.92
150.31
159.16
168.18
Total, no water
65.55
8.21
17.64
Impurities in stream
30.36 %
0%
5.16 %
10.49 %
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From Table 9.1 it is clear that the design option that makes gasohol directly, doesn’t
produce actual waste water, but pure water as far as Aspen Plus 10 can simulate. Pure
water is a valuable product, which can be sold, but in practice this stream may contain
some contaminations. Therefore the water is not seen as a product in this conceptual
design, but as a waste stream, which flows directly to the sewage system. The
azeotropic distillation by toluene contains the highest percentage of impurities as can
been seen in Table 9.1. The amount of ethanol waste is in each process option clearly
distinct. This ‘non-recovery’ is a result of intrinsic process characteristics and choices
in process and process units. This is dealt with in Paragraph 9.3.
In Chapter 3 is already stated that the contaminated waste water will be transported to
a nearby biological purification plant. It will purify the water in exchange for
payment. This choice of waste water treatment and the basis for the costs are
explained in Paragraph 9.2. The actual costs of this treatment will be summarised in
Chapter 11.
9.2
Biological treatment of waste water
As seen in Table 9.1 the waste water streams contain some amount of hydrocarbons.
These components are harmful for humans and environment. The higher the
concentrations of the hydrocarbons, the more harmful a waste water stream is, so the
more has to be paid to the waste disposal plant. Three options of waste water
treatment were considered:
1. Transport to a biological waste water purification plant for payment.
2. Transports to a waste disposal plant that combusts the water for payment.
3. Combustion of the waste water on the plant site.
The first option is quite expensive because the waste water has high a concentration of
hydrocarbons and payment is based on these concentrations (as will be explained
further on).
The other options are combustion. The costs of combustion, at the own plant site or at
another plant, are not known in spite of attempts to obtain them. The costs that have to
be taken into account are: a furnace, fuel and probably equipment to recover produced
steam. Discussions with experienced chemical engineers resulted into the assumption
that combustion is more expensive than or as expensive as the biological treatment.
Specialised waste purification plants exist (in the Netherlands for example the AVR in
Rotterdam), so it is probably cheaper to transport the waste water than to combust it
on the plant site itself, also because of the relatively small stream.
The waste water treatment will be done at a nearby biological purification plant for
payment. The following equations show the conversion of waste water into a quantity
that will express the harm to humans; the Population Equivalent (PE) in number of
people This quantity determines the costs of the contracted out waste water treatment.
These costs are set to 100 guilders per PE per year (€ 45.4 per PE per year). This is
based on ref. 27, p IIA-56
This PE is based on the amount of oxygen necessary to oxidise the domestic waste
water discharged by one person per day and is defined as:
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PE 
(COD  N Kj )  V
in which:
COD
NKj
V
0.136
(9.1)
0.136
Total Chemical Oxygen Demand (kg O2 / m3waste water)
Concentration of nitrogen, determined using the Kjeldahl method (kg
N / m3waste water )
Volume flow of waste water (m3/day)
Average amount of oxygen-using substances (kg O2 / (day . person))
The total COD is the sum of the individual COD’s for every hydrocarbon in the waste
water. In Appendix 13 an example calculation for a COD is done, from where it is
clear that the COD increases when the concentration of a hydrocarbon increases.
Although the Chemical Oxygen Demand is calculated instead of the Biological
Oxygen Demand, the waste water still is treated in a biological waste water
purification plant (ref. 27, p.IIA-56). Because no nitrogen containing components are
present, NKj is zero for all design alternatives. In Table 9.2 a summary of the
toxicological properties of the different waste water streams is given.
Table 9.2: Toxicological properties of waste water streams
Toxicological
Azeotropic
Extractive distillation
property
distillation by
by PAA
toluene
5.56
3.88
V (m3/d)
3
COD (kg O2 / m )
651
131
PE (people)
26,618
3,732
Normal distillation followed
by membrane purification
4.15
234
7,144
From this table it can be seen that the waste water from azeotropic distillation by
toluene has the highest PE and so is the most expensive to be treated in the biological
purification plant. Extractive distillation by PAA appears to be the cheapest in the
field of waste treatment. This is in accordance with the last row of Table 9.1 that
shows the total stream of hydrocarbons.
9.3
Influence of process on wastes
In Paragraph 9.1 is already seen that the four alternatives produce different waste
water streams. This difference is a result of a combination of intrinsic thermodynamic
interaction between the substances and the process and process unit choices.
Therefore equipment often has to be chosen by taking into account the
thermodynamic properties of the relevant mixtures.
From Table 9.1 some thermodynamic interaction in process units can be retrieved.
Because gasoline and water are immiscible the gasohol production plant has a very
pure waste water stream. In the distillation columns using PAA and the zeolite
membrane-unit all the isobutyl and isopentyl alcohol go along with the water into the
waste stream. This is because of the polar (hydrophilic) hydrogen bond between these
alcohols and water. Because ethyl acetate and acetaldehyde aren’t hydrophilic but
hydrophobic substances they have a strong tendency to go along with the ethanol
product. In the ‘standard’ case with toluene as entrainer all four contaminations end
-64-
up entirely in the waste water, mainly by the influence of the specification of ethanol
to be reached (99.8 vol%). Besides the contaminations obviously the recovery of
ethanol must be reviewed when designing.
Besides the influence of inherent thermodynamic relations between the substances,
also the units of equipment used have great influence. Adding some unit operations,
contaminations can be removed from the product. But the correlated costs have also to
be taken into account. The influence of thermodynamic interaction and the process
(unit) choices are explicated here for every design alternative.
9.3.1 Azeotropic distillation by toluene
In this design three distillation columns and one decanter are the main equipment
units. The distillation columns will separate the components to some extent, but some
impurities always remain because of thermodynamic interaction / properties
mentioned above. Separation costs will grow when performance is improved. In this
case three columns cope with this problem and therefore equipment costs can grow
exponentially. The decanter performance is dependent on the operating pressure and
temperature, but it is quite constant because of the interactions between the
substances. It appears this process uses only equipment, which is not able to produce
by definition pure streams, so a choice has to be made between equipment size,
number of trays and columns and waste composition. Because of the feed
contaminations and toluene much ethanol ends up in the waste water stream, what
results in a recovery of 96 w%.
9.3.2 Extractive distillation by gasoline
In this alternative the distillation column obviously produces a contaminated water
stream. But a decanter can separate the water well, because of the difference in
polarity between water and the hydrocarbons (mentioned above).
This process has the advantage that thermodynamic properties of the relevant
mixtures are such that simple process units can be used. As a result the ethanol has a
recovery of 100 w%.
9.3.3 Extractive distillation by PAA
The interaction between ethanol and water are changed here in such way that the
distillation column can produce almost pure ethanol. The bottom stream is slightly
contaminated water containing the polymer PAA. This polymer can easily be
removed from the waste water stream by an ultrafiltration unit. This process has the
advantage that thermodynamic properties are altered and therefore simple equipment
can be used, while still relatively pure waste water will be produced. The recovery of
ethanol reached is 99.9 %.
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9.3.4 Normal distillation followed by membrane purification
The distillation column used has the problem that ethanol can’t be purified further
than the azeotrope. But therefore a membrane-unit is used to selectively remove water
from that stream. This implies that the only real waste is the stream at the bottom of
the distillation column. The result is a ethanol recovery of 99.0 w%.
Summarised, by choosing equipment (possibly in combination with an entrainer) that
takes advantage of the thermodynamic interaction between the substances in the feed
mixture, a relatively clean waste water stream can be produced and costs are
minimised. Obviously costs of the equipment a major issue in equipment choice.
All options contain one or more distillation columns that will maintain contaminations
in the waste water stream. When this stream has to be purified more, the column has
to be changed and becomes more expensive. Especially the extra utility costs in
condenser and reboiler are very high. Probably the column will increase largely in the
number of trays too. Therefore the distillation column can be economically combined
with another piece of equipment or an extra substance.
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10 Process safety
All manufacturing processes are to some extent hazardous, but in chemical processes
there are additional, special hazards present associated with the chemicals used and
the processes conditions operated at. There are several methods available to identify
and classify hazards. Two of them are carried out in this chapter, namely the Dow Fire
and Explosion Index and the Hazard and Operability Studies (ref. 27 and ref.33,
Chapter 9).
10.1 The Dow Fire and Explosion Index
The Dow Fire and Explosion Index is a hazard classification guide that gives a
method to evaluate the potential risk from a process by combining chemical and
engineering factors. A numerical ‘Fire and Explosion Index’ (F&EI) is calculated for
all four plant designs, based on the nature of each process and the properties of the
process materials (see Appendix 20). The larger the value of the F&EI, the more
hazardous the selected process option.
The basis of the F&EI is the Material Factor (MF). The Material Factor is a measure
of the intrinsic rate of potential energy release from the burning, explosion or other
chemical reaction of the material. In calculating the F&EI for a process the value for
the material with the highest MF, which is present in significant quantities is used.
In the azeotropic distillation by toluene the materials present in significant quantities
are ethanol, toluene and water. The ethanol and toluene both have the same value for
the MF. So to make a material choice the health aspects of both materials are taken
into account. In this process option the entrainer toluene is the material chosen in the
calculation of the F&EI, because it is more harmful than the relatively innocent
ethanol. In the extractive distillation by gasoline the candidates present for the MF are
ethanol and gasoline. On the basis of the same principle as in the azeotropic
distillation by toluene, gasoline is chosen for the calculation of the F&EI. In the
extractive distillation by PAA and the normal distillation followed by membrane
purification only ethanol and water are available in significant quantities. So the
material with the highest MF is inevitable ethanol.
The MF is multiplied by a Unit Hazard Factor (F3) to determine the F&EI for the
process unit. The Unit Hazard Factor is the product of two other factors, which take
account of the inherent hazards in the operation of the particular process unit: the
General and Special Process Hazards.
The general process hazards are factors that determine the magnitude of the loss
caused by an incident. Six factors are listed on the calculation form and in Table 10.1.
The calculated values can be seen in Appendix 20.
Table 10.1: General process hazards.
A
Exothermic chemical reactions
B
Endothermic processes
C
Materials handling and transfer
D
E
F
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Enclosed or indoor process units
Access
Drainage and spill control
The Special process hazards are factors that are known from experience to contribute
to the probability of an incident involving loss. Twelve factors are listed on the
calculation form and in Table 10.2. The calculated values can be seen in Appendix 20.
Table 10.2: Special process hazards.
A
Toxic material(s)
G
B
C
D
E
F
H
I
J
K
L
Sub-Atmospheric pressure
Operation in or near flammable range
Dust explosion
Pressure
Low temperature
Quantity of flammable/unstable
material
Corrosion and erosion
Leakage- Joints and packing
Use of fired equipment
Hot oil heat exchange system
Rotating equipment
The Material Factors of the four different process options are obtained from NH, NF
and NR. The NH, NF and NR are ratings expressing health (related to toxicity),
flammability and reactivity (or instability) respectively. Generally NF and NR are for
ambient temperatures. Because the maximum temperatures in all four process options
are over 60 ºC, a temperature adjustment is required for NF. The Temperature
Adjusted Material Factors are listed below in Table 10.3.
Table 10.3: Material Factors for the four process options.
Process
Basic material for MF NH
Azeotropic distillation
by toluene
Toluene
2
Extractive distillation
by gasoline
Gasoline
1
Extractive distillation
by PAA
Ethanol
0
Normal distillation
followed by
membrane purification
Ethanol
0
NF
NR
Temperature adjusted MF
3+1
0
21
3+1
0
21
3+1
0
21
3+1
0
21
Table 10.1 shows that the materials used in all four processes are non-reactive, even
under fire (NR = 0), but exhibit a maximum hazard with regard to flammability (NF =
4). The health factor for the three different basic materials varies. Toluene has the
highest health rating (NH = 2), which means that medical attention is required when
working with this substance to avoid temporary or residual injury. The health rating of
gasoline (NH = 1) stands for the likelihood of only minor injuries. Ethanol imposes no
health hazard towards operating personnel (NH = 0). As far as the material properties
are concerned, there can only be made a distinction between toluene, gasoline and
ethanol on the basis of NH. The final conclusion that can be drawn is that ethanol
exhibits the least degree of hazard on the basis of health.
Table 10.4 is a listing of the calculated F&EI values versus a description of the degree
of hazard to give some relative idea of the severity of the F&EI.
-68-
Table 10.4: Assessment of hazard.
Process
Azeotropic distillation by toluene
Extractive distillation by gasoline
Extractive distillation by PAA
Normal distillation followed by
membrane purification
Fire and Explosion Index
95
113
76
Degree of hazard
Moderate
Intermediate
Moderate
90
Moderate
Table 10.4 indicates that the probability and potential magnitude of a fuel or energy
release resulting from process control failures, equipment failures or from vibration or
other sources of stress fatigue, is the highest for the extractive distillation by a
gasoline entrainer. The qualification intermediate is mainly caused by the presence of
gasoline, which forms a flammable mixture in the presence of air. Extractive
distillation by polyacrylic acid has the lowest Fire and Explosion Index and is
therefore the less hazardous process among the four process options reviewed.
Azeotropic distillation by toluene and normal distillation followed by membrane
purification exhibit about the same degree of hazard.
10.2 Hazard and Operability Studies
A Hazard and Operability Study (HAZOP) is a procedure for the systematic, critical,
examination of the operability of a process. Applied to a process design, it indicates
potential hazards that may arise from deviations from the intended design conditions.
A limited HAZOP study is carried out for evidently critical pieces of equipment in
each process option, using ‘guide words’ to help generate thought about the way
deviations from the intended operating conditions can cause hazardous situations. The
following six guide words are used (see Appendix 21):
- Not, no
- More
- Less
- As well as
- Part of
- Other than
The selected pieces of equipment, on which the HAZOP technique is performed, are
the distillation columns of each design and the membrane unit of the normal
distillation column followed by membrane purification. The results are shown in
Appendix 21. These pieces of equipment are seen as the most dangerous. This is
because they are not operated at ambient temperature and pressure.
The recommended actions to deal with the potential hazards discovered are mainly
modifications to the control system and instrumentation. In other words improve
control, include additional valves and alarms, install spare pieces of equipment and
introduce fire protection of the equipment. Most of the recommended actions are
already implemented in the designs (see Appendices 8 and 15 till 18). Besides the
controllers mentioned in Chapter 6 the designs include a spare pump for each pump in
the processes. For safe operation also a pressure relief valve is installed at the top of
the columns. To prevent line fracture and column heating or fracture thermal
expansion relieves are installed in the lines. The safety measures kickback over feed
-69-
pumps and alternative ways to condense or reboil are not implemented in the design
options. These measures are omitted because of practical and economical reasons.
The actions still to be taken in a real life situation are the introduction of fireextinguishing systems, good communication between operating personnel and the
institution of regular patrolling and inspection of the plant. This is not dealt with in a
conceptual design.
-70-
11 Economy
Continuity is the purpose of every chemical plant, so estimates of the expected profit,
the investment required and the cost of production are needed before the profitability
of a project can be assessed. In this chapter the various components that make up the
capital costs of a plant and the components of the operating costs are discussed. This
chapter is based on the principles used in Coulson and Richardson’s Chemical
Engineering, Volume 6, Chapter 6, Costing and Project Evaluation (ref. 33).
11.1 Capital investment
Capital investment is the sum of the fixed capital and the working capital needed for a
project. Both components of the total investment are discussed below.
The fixed capital is the total cost of the plant ready for start-up. It is the cost paid to
the contractors. It includes the cost of:
1. Design, and other engineering and construction supervision.
2. All items of equipment and their installation.
3. All piping, instrumentation and control systems.
4. Buildings and structures.
5. Auxiliary facilities, such as utilities, civil and land engineering work.
It is a once-only cost that is not recovered at the end of the project life, other than the
rest-value. It is assumed that at the end of the project life, 5 % of the investment costs
is the scrap value.
The working capital is the additional investment needed, besides the fixed capital, to
start up the plant and to operate it to the point when income is earned. The working
capital includes the following costs:
1. Start-up
2. Initial entrainer charges
3. Raw materials and intermediates in the process
4. Finished product inventories
5. Funds to cover outstanding accounts from customers
Most of the working capital is recovered at the end of the project. Working capital can
vary from 5 per cent of the fixed capital for simple, single-product processes; to as
high as 30 per cent for a process producing a diverse range of product grades for a
refined market. Because the four design alternatives are simple, straightforward
processes, the working capital is set to the value of 5.3 % of the fixed capital.
The fixed capital cost for the process options are based on an estimate of the purchase
cost of the major equipment items required for the process, the other costs being
estimated as factors of the equipment cost. In the factorial method of cost estimation
the fixed capital cost of the project is given as a function of the total purchase
equipment cost by the following equation:
Cf  f L  Ce
(11.1)
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in which:
Cf
fL
Ce
Fixed capital cost
‘Lang factor’
Total delivered cost of all major equipment items
The cost factors that are compounded into the ‘Lang factor’ are considered
individually. The direct-cost items that are incurred in the construction of a plant, in
addition to the cost of equipment are listed below in Table 11.1 (f1 through f9). In
addition to the direct cost of the purchase and installation of equipment, the capital
cost of a project will also include the indirect costs listed in Table 11.1 (f10 through
f12). These can be estimated as a function of the direct costs.
Table 11.1: Factors for the estimation of the process options fixed capital cost.
Item
Symbol
Factor
1. Major equipment, total purchase cost
PCE
Equipment erection
0.40
f1
Piping
0.70
f2
Instrumentation
0.20
f3
Electrical
0.10
f4
Process buildings and structures
0.15
f5
Utilities
f6
Storages
0.15
f7
Site development
0.05
f8
Ancillary buildings
0.15
f9
2. Total physical plant cost (PPC)
PPC
Design and Engineering
0.30
f10
Contractor’s fee
0.05
f11
Contingency
0.10
f12
The contribution of each of these items to the fixed capital cost is calculated by
multiplying the total purchased equipment by the appropriate factor (equation 11.2
and 11.3).
PPC  PCE  (1  f1  ...  f9 )
(11.2)
FCC  PPC  (1  f10  f11  f12 )
(11.3)
in which:
FCC
PPC
PCE
fa(b)
Fixed capital cost
Total physical plant cost
Purchase cost equipment
Lang factor
A summary of the purchased equipment costs (ref. 38) and the total investment
needed for the four different process options is shown in Table 11.2. The calculations
of the purchase cost of major equipment items, total physical plant costs, fixed capital,
working capital and total investments required for the four alternatives are listed in
Appendix 22.
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Table 11.2: Purchase cost of equipment and total investment for all four processes.
Process
Distillation
Heat
Pumps Vessels
Ultrafiltration
Total
option
columns
exchangers
unit
Investment
(k€)
(k€)
(k€)
(k€)
(k€)
(k€)
Azeotropic
distillation
by toluene
262
100
42
15
1,859
Extractive
distillation
by gasoline
41
87
26
10
730
Extractive
distillation
by PAA
116
85
25
6
25
1,151
Distillation
followed by
membrane
13
24
103
68
purification
930
From Table 11.2 it can be seen that the azeotropic distillation by toluene requires the
highest total investment. This is caused by the presence of the three distillation
columns in the design. Extractive distillation by gasoline is the process option that
needs the lowest total investment. The reason for this is the relatively small
distillation column used. From the table it can be concluded that the purchase cost of
the distillation columns is the decisive factor for the height of the total investment.
When these fixed capital costs are compared with the maximum allowable investment
calculated in Chapter 3.4 it is notable that the values in Chapter 3 are much lower than
the values mentioned above.
11.2 Operating costs
To judge the viability of the four alternatives and to make choices between them an
estimate of the operating costs, the costs of producing the product, is needed. The
costs of producing a chemical product are divided into three groups, which are listed
in Table 11.3 (see Appendix 22):
Table 11.3: Operating costs.
Variable operating costs:
1. Raw materials
2. Miscellaneous operating materials
3. Utilities
4. Shipping and packaging
10 % of maintenance costs
Negligible
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Table 11.3 (continued)
Fixed operating costs:
5. Maintenance
6. Operating labour
7. Laboratory
8. Supervision
9. Plant overheads
10. Capital charges
11. Insurance
12. Local taxes
13. Royalties
General operating expenses:
14. Sales expenses
15. General overhead
16. Research and development
5 % of fixed capital
20 % of operating labour
20 % of operating labour
50 % of operating labour
10 % of fixed capital
1 % of fixed capital
Left out of consideration
Left out of consideration
15% of variable and fixed operating costs
In these designs the shipping and packaging costs are assumed to be negligible. For
comparing economic competitiveness of the processes, in first instance local effects as
local taxes and royalties are left out of consideration. The operating labour consists of
one operator each of five shifts. The annual salary of an operator is assumed to be €
40,000.
For the purchase of raw materials the prices are based on several literature data. These
data are summarised in Table 11.4 together with the costs of the utilities.
Table 11.4: Costs raw materials and utilities.
Raw materials:
Unity
Ethanol 88 w% 1
ton
Gasoline2
ton
2
Toluene
ton
Polyacrylic acid 1
ton
EUR per unity
250
1200
450
1300
Utilities:
Cooling water 3
MP steam 3
Electricity 3
Nitrogen 4
EUR per unity
450.00
12.50
50.00
1.35
Unity
kton
ton
MWh
ton
1 Estimation
2 ref. 10
3 Appendix 3
4 ref. 38
For the ethanol feed stream no prices are available. Therefore a rough estimation is
made of approximately 45 % of the sales price of pure ethanol. This percentage is
chosen to express the inlinearity of the ethanol price to its purity. A pure ethanol
stream can be used for much more applications and is therefore much more valuable
than an aqueous, diluted ethanol stream like the feed stream. Most of the costs of
purifying the ethanol are energy costs and the investment costs to overcome the
azeotrope. Therefore the sales price is estimated to be more than twice the price of an
aqueous stream. For the polyacrylic acid a price of EUR 1300/ton is estimated on
basis of common prices for ordinary polymers.
-74-
The total production costs of each alternative are listed in Table 11.5. The extended
calculation can be found in Appendix 22.
Table 11.5: Manufacturing costs provided in amount of money per year and per ton of
product.
Process option
Total production costs
Total production costs
(k€/year)
(€/ton product)
Azeotropic distillation by
toluene
4,300
518
Extractive distillation by
gasoline
111,540
1,286
Extractive distillation by
PAA
4,325
497
Normal distillation followed
by membrane purification
4,133
481
The extractive distillation by gasoline has the highest annual operating costs due to
the high amounts of gasoline needed to manufacture the desired gasohol product. The
other three process alternatives do not differ very much from each other in
manufacturing costs.
11.3 Income
The gross income of a project is the sum of the product sales minus the costs of the
disposal of the waste water (see Chapter 9 – Wastes).
The sales prices of ethanol and gasohol and the costs of waste water are listed in
Table 11.6.
Table 11.6: Sales prices of the products and costs of waste water disposal.
Materials
Ethanol 99,8 w% 1
Gasohol 2
Waste water 3
Unity
ton
ton
PE
EUR per unity
550
1296
- 45
1 ref. 10
2 assumption
3 Appendix 13
For the price of gasohol an assumption is made. Assumed that the gasohol prices are
higher than the equivalent of ethanol and gasoline a factor of 14 % grant is added. The
local government or European Parliament should provide this grant to encourage the
use of biofuel. The sensitivity of the amount of grant will be discussed in Paragraph
11.7.
The gross annual income of the four process options is summarised in Table 11.7.
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Table 11.7: Gross Annual Income.
Process option
Sales of product
(k€/year)
Azeotropic distillation
by toluene
4,565
Extractive distillation
by gasoline
112,428
Extractive distillation
by PAA
4,785
Distillation followed by
membrane purification
4,730
Costs of waste water disposal
(k€/year)
Gross Income
(k€/year)
-1,208
3,357
-
112,428
-169
4,616
-324
4,406
From the table it can be seen that the extractive distillation by gasoline results in the
highest annual income due to the gross profit of the sale of gasoline and the absence
of waste water treatment costs. Azeotropic distillation has the lowest gross income a
year as a result of the high costs of waste water disposal.
11.4 Cash flow
The flow of cash is essential for commercial organisations. The ‘net cash flow’ at any
time is the difference between income and operating costs. The cash flows are based
on the estimates of investment, operating costs, sales volume and sales price that are
made for each design alternative.
In Appendix 22 three different cash flows are calculated. The first one is the net cash
flow for the project considered as an isolated system, and taxes on profits and the
effect of depreciation of the investment are not considered. The second one is the net
cash flow after annual depreciation of the total investment over fifteen years. The last
one is the net cash flow after depreciation and tax. The tax rate is set at 40 % for
positive cash flows and at 0 % for negative cash flows. The cash flows are lined up in
Table 11.8.
Table 11.8: Cash Flows of the four different process options.
Process option
Net Cash Flow,
Net Cash Flow,
before tax
after depreciation
(k€)
(k€)
Azeotropic distillation by
toluene
-943
-1,066
Extractive distillation by
gasoline
888
840
Extractive distillation by
PAA
290
214
Normal distillation followed
by membrane purification
273
211
Net Cash Flow,
after tax
(k€)
-1,066
504
128
127
The cash flow of the azeotropic distillation by toluene has a negative value, so no
profit is made in operating this process option.
-76-
11.5 Economic criteria
As the purpose of investing money in chemical plants is to earn money, some means
of comparing the economic performance of the four process options is needed. The
criteria used to judge their economic performance are: the Pay-Out Time, the Rate on
Return, the Discounted Cash Flow Rate on Return, the Net Present Value and the Net
Future Value. These criteria provide the link between ‘Once-Off’ investment and
annual income and costs and are explained in this paragraph. At the end of the
paragraph a summary of all the comparing methods will be made for all four options.
11.5.1 The Pay-Out Time
The Pay-Out Time (POT in years) is the time required after the start of the projects to
pay off the initial investment from income. This method is not a measure of the
profitability, but is a measure to investigate when the break-even point is reached.
11.5.2 The Rate on Return
The Rate on Return (RoR in %/year) is the ratio of annual profit to investment
(equation 11.4). It is a simple index of the performance of the money invested. The
Rate on Return is related to the Pay-Out Time as shown in equation 11.5.
RoR  100% 
RoR 
Net Cash Flow
Total Investment
(11.4)
100 %
POT
(11.5)
11.5.3 Net Present Value and Net Future Value
The money earned in any year of the four projects can be reinvested as soon as it is
available and start to earn a return. So money earned in the early years of the project
is more valuable than that earned in later years. This ‘time value of money’ can be
allowed for by using equations (11.6) and (11.7). The net cash flow in each year of the
projects is brought to its ‘present value’ at the start of the project by discounting it at
some chosen compound interest rate (r).
NPVnow, year n 
n t
NPVtotal  
n 1
in which:
NPVnow,year n
NFVyear n
(11.6)
(1  r ) n
NFVyear n
(11.7)
(1  r ) n
Net Present Value in year n brought back to present
value (€)
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NFVyear n
r
n
NPVtotal
t
Net Future Value in year n (€)
Compound interest rate (-)
Time after project start (years)
Total Net Present Value of project (€)
Life of project (years)
The interest rate is chosen to reflect the earning power of money. It would be roughly
equivalent to the current interest rate that the money could earn if invested. An
interest rate of 10 % is taken. This rate is taken on basis of the European interest rates
of the last 10 years. This rate fluctuates between 10.08 % (1992 Q3) and 3.99 % (1999
Q1). At the moment the interest rate is approximately 4,7% (ref. 23). To ensure that
the taken interest rate is sufficient and because it is difficult to predict the future rates,
the highest interest rate of the last 10 years is taken.
The total NPV is less than the total NFV, and reflects the time value of money and the
pattern of earnings over the life of the projects.
11.5.4 Discounted Cash Flow Rate on Return
Discounted cash-flow analysis, used to calculate the present value of future earnings,
is sensitive to the interest rate assumed. By calculating the NPV for various interest
rates, it is possible to find an interest rate at which the cumulative net present value at
the end of the project is zero. This particular rate is called the Discounted Cash Flow
Rate on Return (DCFRoR) and is a measure of the maximum rate that the project
could pay and still break even by the end of the project life (equation 11.8).
n t
NFVyear n
 (1  r ')
n 1
n
0
in which:
r’
NFVyear n
(11.8)
Discounted cash flow rate on return (DCFRoR) (-)
Future value of the net cash flow in year n (€)
The value of r’ is found by trial-and-error calculations. Finding the discount rate that
just pays off the project investment over the project’s life is analogous to paying off a
mortgage. The more profitable the project, the higher the DCFRoR that it can afford
to pay.
11.5.5 Summary
The economic criteria discussed in this section are set out for the four process
alternatives in Table 11.9.
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Table 11.9: Economic criteria.
Process option
POT
(years)
Azeotropic distillation
by toluene
-2.0
Extractive distillation
by gasoline
0.8
Extractive distillation
by PAA
4.0
Normal distillation
followed by
membrane purification
3.4
RoR
(%)
DCFRoR
(%)
NPV
(k€)
NFV
(k€)
-50.7
-
-7,839
-14,019
121.6
85.2
5,048
10,855
25.2
19.1
790
2,681
29.4
24.6
886
2,664
As can be seen Table 11.9 the azeotropic distillation is unprofitable due to its negative
cash flow. The investment costs will be never earned back. The other options are all
profitable. The forecast on basis of these economic criteria is for the extractive
distillation by gasoline most positive. This option has its break even point already
after 0.8 years and has the highest net future value and the highest net present value.
11.6 Cost review
In this paragraph an indication will be given of the main costs elements. In Appendix
22 the various costs are calculated as percentage of the total production costs. In Table
11.10 a summary is made of the three main costs for each option.
Table 11.10: Summary of the main costs as percentage of the total production costs.
Process option
Raw materials
General overhead
Utilities
Azeotropic distillation by toluene
60 %
13 %
12 %
Extractive distillation by gasoline
86 %
13 %
0.3 %
Extractive distillation by PAA
58 %
13 %
16 %
Normal distillation followed by
membrane purification
61 %
13 %
13 %
Table 11.10 shows that the main costs of production are the purchase of the raw
materials. The costs of the purchase are about 60 % of the total production costs
except extractive distillation by gasoline where the costs of raw materials amount to
86 %. Other important costs are the overhead costs and the costs of the utilities. When
a reduction in the costs is necessary or desirable the main costs should be taken in
consideration first. For example a cost reduction in purchase of the raw materials can
be made by making discount agreements with suppliers. Another possibility of
reducing the cost can be for example further integration and optimisation of the
utilities.
11.7 Sensitivities
Because the price of ethanol is predicted to decrease and the price of the feedstocks to
increase (Window, Chapter 1), it is of great importance to investigate fluctuations in
the cost of the raw materials. To investigate the influence of small changes in the
costs of raw materials, a calculation is made by taking the raw material costs 5 %
-79-
higher. The results are shown in Table 11.11. For the complementation the same is
done for in case the raw materials will be decrease 5 % in costs in Table 11.12. The
relative influence will be shown as percentages of the previous data shown in
Paragraph 11.5.5.
Table 11.11: Absolute and relative influence on economic criteria by changing the prices of
the raw materials (-5%).
Economic
Azeotropic
Extractive
Extractive
Normal
criteria
distillation by
distillation by
distillation by
distillation
toluene
gasoline
PAA
followed by
membrane
purification
POT (years)
RoR (%)
DCFRoR (%)
NPV (k€)
NFV (k€)
Abs
-2.3
-42.8
-6,888
-12,104
Rel (%)
+15
-16
-12
-14
Abs
0.1
878.3
330.9
40,736
82,699
Rel (%)
-88
+929
+622
+707
+662
Abs
2.6
37.8
30.5
1,723
4,560
Rel (%)
-35
+50
+60
+118
+70
Abs
2.2
44.9
37.2
1,819
4,542
Rel (%)
-35
+53
+51
+105
+70
Table 11.12: Absolute and relative influences on economic criteria by changing price raw
materials (+5%).
Economic
Azeotropic
Extractive
Extractive
Normal distillation
criteria
distillation by
distillation by
distillation by
followed by
toluene
gasoline
PAA
membrane
purification
POT (years)
RoR (%)
DCFRoR (%)
NPV (k€)
NFV (k€)
Abs
-1.7
-58.6
-8,790
-15,934
Rel (%)
-15
+16
+12
+14
Abs
-0.2
-635.1
-30,639
-60,989
Rel (%)
-125
-622
-707
-662
Abs
7.9
12.7
4.7
-143
803
Rel (%)
+98
-50
-75
-118
-70
Abs
7.2
13.8
9.0
-47
785
Rel (%)
+112
-53
-63
+105
-71
As can be seen in Tables 11.11 and 11.12 a both a positive as negative change in the
price of the raw materials have huge influence on the economic criteria. Especially the
process extractive distillation by gasoline is extremely sensitive to fluctuations in
price. When the purchase prices slowly rise, the profits of the gasohol production are
towering, while small increases in the prices the gasoline unit showered under dept.
The extractive distillation by PAA and the normal distillation followed by membrane
purification are also quite sensitive for price changes. Even if a the raw materials are 5
% more expensive, profit can still be made under the restriction of interest rates are
not higher than respectively 4.7 % and 9.0 % The azeotropic distillation by toluene is
even at low cost prices of its raw materials not profitable.
Also the sensitivity for the total fixed costs will be investigated. The fixed costs will
be varied by –10 % and 10 %. In Table 11.13 the absolute and relative changes will be
shown.
-80-
Table 11.13: Absolute and relative influences on economic criteria by changing fixed
capital by +10% and –10%.
Fixed
Azeotropic
Extractive
Extractive
Normal
capital
distillation by
distillation by
distillation by
distillation
toluene
gasoline
PAA
followed by
membrane
purification
Abs
POT
(years)
RoR
(%)
DCFRoR
(%)
NPV
(k€)
NFV
(k€)
+10%
-10%
+10%
-10%
+10%
-10%
+10%
-10%
+10%
-10%
-2.1
-1.8
-47.7
-54.3
-8,231
-7,447
-14,632
-13,407
Rel
(%)
+5
-10
-6
+7
+5
-5
+4
-4
Abs
0.9
0.7
108.9
137.2
78.2
93.4
4,895
5,202
10,615
11,095
Rel (%)
+13
-13
-10
+13
-8
+10
-3
+3
-2
+2
Abs
4.7
3.3
21.3
30.0
15.1
23.7
548
1,033
2,302
3,060
Rel (%)
+18
-18
-15
+19
-21
+24
-31
+31
-14
+14
Abs
4.0
2.9
25.0
34.6
20.8
29.1
690
1,081
2,357
2,970
Rel (%)
+18
-15
-15
+18
-15
+18
-22
+22
-12
+11
In the table can be seen that fluctuations in the fixed capital have less influence in the
economic criteria than fluctuations in the purchase price. However, a change of 10 %
in the fixed costs still influences the economic criteria often with more than 10 %.
11.8 Negative cash flows
The option azeotropic distillation seems to be unprofitable due to its negative Cash
Flow. The production costs are higher than the income from the sales. As is shown in
Paragraph 11.6 the main costs in the production costs are the raw materials, general
overhead and the utilities. The raw materials are variable costs and may be purchased
cheaper by agreements with suppliers. It can be calculated that a discount of more
than 36 % in the raw materials should be attained to break even within 15 years. The
prices of the general overhead and the utilities are fixed by respectively a fixed
percentage of the operating costs and the assignment.
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12 Comparison and conclusions
This report covers the design of four ethanol dehydration processes, which produce
ethanol to be used as a fuel additive. From the abundant literature, four dehydration
possibilities are chosen on the basis of the use of modern techniques, the use of nonhazardous materials, economical profitability, data availability and personal interests.
After consideration a choice is made for:
- Azeotropic distillation by toluene
- Extractive distillation by gasoline
- Extractive distillation by polyacrylic acid (PAA)
- Normal distillation followed by membrane purification
The purpose of this report is to make a comparison on the basis of data validity, purity
and recovery, process yields, wastes, process safety and economy between the four
possibilities of ethanol dehydration and to make a recommendation for a further
design of one of the process options.
Ethanol is not suitable for the use as fuel additive until it is dehydrated and purified
till 99.8 vol%. Therefore the designs has to purify the ethanol feed stream of 88.8
vol% by removing the pollutants water, acetaldehyde, ethyl acetate, isobutyl alcohol
and isopentyl alcohol, except for the extractive distillation by gasoline. This design
already complies with the conditions of the gasohol.
12.1 Data validity
The thermodynamic relations and estimation methods used in the process designs
should be valid for the temperatures and pressures occurring in the equipment.
Therefore in Chapter 4 the data validity of the four options is investigated. It appeared
that the UNIQUAC model is the most appropriate thermodynamic model. In the case
of extractive distillation by PAA a problem arises because the available flowsheet
simulation program Aspen Plus 10 cannot simulate polymers. Therefore the vapourliquid equilibrium data from literature are fit in a UNIQUAC model by the maximumlikelihood method. The calculated parameters are however very sensible to pressure
changes and therefore some doubt arises about the validity of the model used.
12.2 Purity and recovery
Each design complies with the criterion of an ethanol product purity of 99.8 vol%. For
the alternative extractive distillation by gasoline this criterion is converted in a
maximally allowed absolute amount of water, namely 5.6.10-4 kg/s, in the produced
gasohol stream. With a value of 1.8.10-4 kg/s, this criterion is amply achieved. The
azeotropic distillation by toluene has the highest purity with 99.95 vol%. The
recovery of the plants differ from 95.5 % (azeotropic distillation by toluene) to 99.9 %
(extractive distillation by PAA). The results are summarised in the Table 12.1.
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Table 12.1: Achieved purities and recoveries.
Process option
Purity
(vol%)
Azeotropic distillation by toluene
99.95
Extractive distillation by gasoline
Extractive distillation by PAA
99.89
Normal distillation followed by
membrane purification
99.86
Purity
(w%)
99.94
99.88
Recovery
(kg/kg)
0.955
1.000
0.999
99.84
0.990
12.3 Process yields
The process yields are calculated for the feedstocks, process chemicals, wastes and
utilities. As can be seen in Table 12.2 the azeotropic distillation by toluene has a high
consumption of feedstock and therefore a high production of waste water compared
with the other options. The consumption of the utilities is in the same order of
magnitude for all alternatives.
Table 12.2: Process yields of the designs
Azeotropic
distillation
by toluene*
Ethanol feed
Waste water
Process chemicals
Cooling water
Electricity
Steam
t/t product
t/t product
t/t product
t/t product
kWh/t product
t/t product
1.21
0.22
0.01
61.6
1.12
2.60
Extractive
distillation
by
gasoline*
1.16
0.15
8.99
44.3
2.48
1.94
Extractive
distillation
by PAA*
1.16
0.16
79.9
1.04
3.39
Distillation
followed by
membrane
purification*
1.17
0.17
64.1
0.80
2.82
* on the basis of ethanol
12.4 Wastes
The main waste is the continuous waste water stream coming out of each designed
plant option. A review of the waste water treatment is used to compare the
environmental impact of the dehydration units. The waste water stream is converted
into Population Equivalent (PE), which is shown in Table 12.3. The costs of a
(biological) waste water stream treatment increase when the population equivalent
increases.
Table 12.3: Population equivalents
Design
Azeotropic distillation by toluene
Extractive distillation by gasoline
Extractive distillation by PAA
Normal distillation followed by membrane purification
Population equivalents (people)
26,618
0
3,732
7,144
As is shown in the table the azeotropic distillation by toluene has a high value of the
PE and is most harmful for the environment, due to the high percentage impurities
ethanol and toluene in the stream. Therefore the treatment costs of this option are the
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highest of the four alternatives. The waste water of the extractive distillation by
gasoline contains merely water, so this stream can be drained away into the sewage.
12.5 Process Safety
The Dow Fire and Explosion Index and a HAZOP study for the critical pieces of
equipment are used to compare the dehydration alternatives as far as safety is
concerned. The Dow Fire and Explosion Index can be used to evaluate the potential
risk from a process by combining chemical and engineering factors. A numerical ‘Fire
and Explosion Index’ (F&EI) and its accompanying degree of hazard is calculated for
all four plant designs and is shown in Table 12.4.
Table 12.4: Results of the Fire and Explosion Index.
Process
Fire and Explosion Index
Azeotropic distillation by toluene
95
Extractive distillation by gasoline
113
Extractive distillation by PAA
76
Normal distillation followed by
membrane purification
90
Degree of hazard
Moderate
Intermediate
Moderate
Moderate
The qualification intermediate for the option extractive distillation by gasoline is
mainly caused by the presence of gasoline, which forms a flammable mixture in the
presence of air. Extractive distillation by polyacrylic acid has the lowest Fire and
Explosion Index and is therefore the less hazardous process among the four process
options reviewed.
The HAZOP study has the same results for all four designs. The suggested controllers,
safety measures and spare pieces of equipment are taken into account in the designs.
On the basis of the performed HAZOP no distinction can be made between the design
options.
12.6 Economy
12.6.1 Economic criteria
For the comparison of the four options in the field of economy, several criteria are
calculated which provide a link between the fixed investment costs and the annual
income and costs. The criteria used to judge their economic performance are: the PayOut Time, the Rate on Return, the Discounted Cash Flow Rate on Return, the Net
Present Value and the Net Future Value. In Table 12.5 the economic criteria are
summarised for the four alternatives.
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Table 12.5: Economic criteria.
Process option
POT
(years)
Azeotropic distillation
by toluene
-2.0
Extractive distillation
by gasoline
0.8
Extractive distillation
by PAA
4.0
Distillation followed
by membrane
purification
3.4
RoR
(%)
DCFRoR
(%)
NPV
(k€)
NFV
(k€)
-50.7
-
-7,839
-14,019
121.6
85.2
5,048
10,855
25.2
19.1
790
2,681
29.4
24.6
886
2,664
As can be seen Table 12.5 the azeotropic distillation is unprofitable due to its negative
cash flow. The investment costs will never be earned back. The other options are all
profitable. The forecast on basis of these economic criteria is most positive for the
extractive distillation by gasoline. This option has its break even point already after
0.8 years and has the highest Net Future Value and the highest Net Present Value.
12.6.2 Sensitivity analysis
Analysing the costs, it appears that the raw materials are a considerable part (58 % to
86 %) of the total production costs. Because the prices of the raw materials are not
fixed a sensitivity analysis is executed to investigate the influence of small price
changes of 5 % on the economic criteria. The relative changes of the criteria are
shown in Table 12.6:
Table 12.6: Relative influence on economic criteria by changing the prices of the raw
materials (-5% and +5%).
Economic
Azeotropic
Extractive
Extractive
Distillation
Criteria
distillation by
distillation by
distillation by PAA
followed by
toluene
gasoline
membrane
purification
Price change
-5 %
+5 %
-5 %
+5 %
-5 %
+5 %
-5 %
+5 %
+15 % -15 %
-88 %
-125 %
-35 %
+98 %
-35 % +112 %
POT
-16 % +16 % +929 %
-622 %
+50 %
-50 %
+53 %
-53 %
RoR
+622 %
+60 %
-75 %
+51 %
-63 %
DCFRoR
-12 % +12 % +707 %
-707 % +118 % -118 % +105 % +105 %
NPV
-14 % +14 % +662 %
-662 %
+70 %
-70 %
+70 %
-71 %
NFV
As can be seen in Table 12.6 both a positive and negative change in the price of the
raw materials have a large influence on the economic criteria. Especially the design
extractive distillation by gasoline is extremely sensitive to fluctuations in price. When
the raw material costs decrease by a fraction, the profits of the gasohol production are
towering. Whereas small increases in the prices showers the gasoline unit under debts.
Also the other options are quite sensible to price changes in the raw materials. The
azeotropic distillation by toluene is still unprofitable, even by lower raw material
prices.
-85-
12.7 Selection for recommendation for further design
To select one of the process options for further design the criteria discussed above are
taken into account. Some criteria however are of such an importance that alternatives
which do not satisfy this criteria are not selected for a final design. The criteria can be
reduced until three most important criteria are left. These criteria are:
- data validity
- profitability
- economical sensitivity
The results of these selection criteria are shown in Table 12.7.
Table 12.7: Results of the comparison on the basis of the main criteria.
Design
Data validity
Economic
criteria
Azeotropic distillation by toluene
+
Extractive distillation by gasoline
+
+
Extractive distillation by PAA
+
Normal distillation followed by membrane
purification
+
+
Economical
sensitivity
+
+
+
As can be seen in Table 12.7, the azeotropic distillation by toluene is rejected on the
basis of the economic criteria (see Paragraph 12.6.1). This option is not profitable and
therefore this option is commercially not interesting. The extractive distillation by
gasoline is rejected on basis of the economical sensitivity (see Paragraph 12.6.2).
Although this process can make an enormous amount of money when the prices in
industry are well, this process also can make enormous debts when small changes in
the purchase of the raw materials occur. The extractive distillation by PAA is rejected
on basis of the doubts that arise concerning the validity of the available data and the
thermodynamic model (see Paragraph 12.1). At this moment not enough data is
available, so before this option could be considered for a final design, more scientific
experiments should be done. In the normal distillation followed by membrane
purification these disadvantages do not occur, so this option is the only robust design.
Ergo the final recommendation of this report is to design the normal distillation
followed by membrane purification in further detail, because this process option is the
most promising of the four.
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13 Recommendations
In this chapter suggestions are done about alternative unit operations, equipment lineup and research areas of uncertainty for the four process alternatives.
13.1 General recommendations
In all process options air-cooled exchangers should be considered. They can be used
both for cooling and condensing. Air-cooled exchangers consist of banks of finned
tubes over which air is blown or drawn by fans mounted below or above the tubes. In
moderate climates air cooling is usually the best choice for minimum process
temperatures above 65 ºC and water cooling for minimum process temperatures below
50 ºC. Between these temperatures a detailed economic analysis is necessary to decide
the best coolant. In all designs except for the extractive distillation by gasoline the
temperature of the streams leaving the top of the columns are above 65 ºC. So aircooled condensers are an interesting option. It is also possible to execute a part of the
regular coolers as air-coolers instead of water-coolers in the various process options.
It can be concluded that the use of air-cooled exchangers can be promising in all four
alternatives. It is clear that air-cooling still needs further investigation. In particular
the Lithuanian climate and the economic profitability have to be examined in detail to
make a final decision between air-cooling and water-cooling.
In all four processes a column-integrated kettle reboiler is used under the assumption
that the costs are competitive or lower than using a thermosyphon reboiler. In practice
a thermosyphon reboiler is the most used reboiler and is considered to be an
economical reboiler. Further investigations and calculations are needed to decide
which reboiler truly is the most economical option to evaporate the bottom streams of
the columns.
In this report the waste water streams are treated in exchange for payment in a
biological water purification plant. Besides biological treatment the waste water can
also be disposed off by combustion. Combustion can take place on the plant site itself
or can be done by a specialised company. The advantage of combustion on the plant is
that steam can be generated from the heat produced during the combustion. This
steam can be utilised in the process heat-exchangers. The costs of combustion, at the
own plant site or at a specialised company, are not known in spite of attempts to
obtain them. It is assumed that combustion is more expensive than or as expensive as
the biological treatment. This assumption has to be checked by concerned companies.
13.2 Azeotropic distillation by toluene
In the azeotropic distillation by toluene it is possible to combine the reflux
accumulator (V01) of column (C01) with the decanter (S01) (see Appendix 8). By
placing a liquid overflow wall in the new combined vessel the two vessels can be
brought together. Behind the overflow wall the toluene is pumped out and transported
to column (C02). At the front site of the overflow wall the ethanol / water stream is
pumped out and divided in a reflux stream to column (C01) and a waste water stream.
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The combination of two relatively small pieces of equipment into one larger vessel
gives a substantial reduction of the costs.
13.3 Extractive distillation by gasoline
In the option extractive distillation by gasoline the possibility exist to combine some
pieces of equipment to operate more economical. From the Process Flow Scheme
(Appendix 8) it can be seen that it is possible to combine the coolers (E01) and (E03)
and maintain the same conditions in stream <11>. The combination of the two coolers
results in a lower amount of cooling water needed and thus in heat integration. The
side effect of this combination is a lower temperature of the reflux stream to column
(C01). This effect can be leveled out by lowering the reflux stream to the column. It is
also possible to combine the reflux accumulator (V01) with the decanter (S01) by
placing an overflow wall. Behind the overflow wall the gasoline is pumped out and
recycled to the column. At the front side the water / ethanol mixture is pumped out
and divided in a reflux stream and a waste water stream. The advantage is the fall out
of (S01), which gives a reduction in costs. The disadvantage is the water / ethanol
waste water stream from (V01) that needs to be purified. It is assumed that the costs
of the purification of this waste water stream are higher than the reduced equipment
costs by joining two vessels together. So the combination of the reflux accumulator
(V01) with the decanter (S01) is not considered as a cost reducing measure.
13.4 Extractive distillation by PAA
The extractive distillation by PAA is a special case. According to literature (ref.2) the
separation of the ethanol / water mixture is accomplished due to the shift of the
azeotropic point by adding the polymer polyacrylic acid. In this article experimental
values for ethanol / water / PAA phase-equilibria are given. The authors found that the
ethanol / water azeotrope disappears when 0.45 w% PAA is added to the mixture.
However, because of the very small amount of experimental data points at high
ethanol mole fractions, which is near the possible azeotrope, it can not be concluded
with certainty that PAA really breaks the azeotrope. Furthermore the pressure
influence on the ethanol / water / PAA phase-equilibria is not yet investigated
sufficiently. The data in the article are only valid at the pressure of 1 bar.
Obviously there is doubt if this article is valid, so the experimental data should be
verified and further data should be obtained before final design and construction can
be realised.
13.5 Normal distillation followed by membrane purification
In the normal distillation followed by membrane purification the reflux pump (P02)
and the product pump (P03) can be combined into one larger pump. The split of the
reflux and the membrane unit feed stream takes place after the new combined pump.
This results in an economical advantage, because one large pump is relatively cheaper
than two small pumps.
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Another very different option to purify the ethanol by membrane purification is the
use of a hydrophobic zeolite membrane (ref. 25). Such a membrane will let ethanol
through as the permeate stream and will let the water flow by. The inlet stream of the
membrane is the aqueous ethanol stream from the ethanol fermentation at 12 vol%. In
this case no distillation column is necessary. It has to be investigated if the gain of
leaving out the distillation column counterbalances the costs of the required size of the
hydrophobic membrane unit. The forecast is that this is a very economical
dehydration option. This alternative is not designed because only alternatives with the
same battery limits can be compared in a consistent manner. In the hydrophobic
membrane option the battery limits are shifted upstream.
-89-
Literature
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2315-2335.
7. Chemiekaarten, Gegevens voor veilig werken met chemicalien, 16de editie, (2001).
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products and semi-conductors, ICISLOR, Reed business information, 2001,
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edition, John Wiley & Sons, New York, (1991), 16, p 186.
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ethanol, Patentnumber US5035776, (1991).
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membranes, Journal of Membrane Science, 144, (1998), p 161-171.
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Text symbols
symbol
description
SI units
A
A
B
C
Ce
Cf
COD
d
D
D
f`
f
fˆ
FCC
g
HETP
L
H
K
n
N
NFV
NFV
NPV
P
PE
PCE
POT
PPC
r
r’
t
RoR
T
U
x
y
UNIQUAC regression parameter
Area
UNIQUAC regression parameter
UNIQUAC regression parameter
Total delivered cost of all major equipment items
Fixed capital cost
Chemical Oxygen Demand
diameter
diameter
UNIQUAC regression parameter
Fugacity
Lang factor
Fugacity of mixture
Fixed capital cost
Acceleration of gravity
Higth of Equivalent Theoretical Plate
Length
Height
Vapour-liquid equilibrium constant
time after project start
Concentration of nitrogen
Net Future Value
Future value of the net cash flow in year n
Net Present Value
Pressure
Population Equivalent
Purchase cost equipment
Pay-Out time
Total physical plant cost
relative compound interest rate
Discounted cash flow rate of return (DCFRoR)
life of project
Rate of return
Temperature
Heat transfer coefficient
Mole fraction in the liquid phase
Mole fraction in the vapour phase
m2
K
€
€
kg O2/m3ww
m
m
K-1
Pa
Pa
€
m/s2
m
m
m
s
kg N/m3ww
€/s
€
€/s
Pa
people
€
s
€
s
%/s
K
W/(m2.K)
-
Greek
description
SI units




ˆ
Activity coefficient
difference
efficiency
Flow
Fugacity coefficient of the vapour phase mixture
m3/s
-
-92-


Density
UNIQUAC binary interaction parameter
kg/m3
-
subscript
description
inner
ij
j
Kj
now,year n
outer
total
tube
V
year n
inner
involving i and j
component j
determined using the Kjeldahl method
in year n brought back to present value
outer
total of a project
tube
volume
n years after start of the project
superscript
description
v
l
vapour
liquid
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