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UNIVERSITY OF GHANA
SCHOOL OF ENGINEERING
PLANT DESIGN FOR THE PRODUCTION OF SORBITOL FROM
CASSAVA
A PROJECT REPORT SUMMITED TO THE DEPARTMENT OF FOOD
PROCESS ENGINEERING
UNIVERSITY OF GHANA, LEGON
BY
ACHARIBASAM VALENTINE 10417033
ANANEY-OBIRI DANIEL 10402033
POKU FRANCIS 10421419
IN PARTIAL FULFILLMENT OF THE REQUIREMENT FOR THE
AWARD OF BSc. ENGINEERING DEGREE IN
FOOD PROCESS ENGINEERING
MAY 2016
Copyright ยฉ2016 University of Ghana
All rights reserved
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DECLARATION
We , Acharibasam Valentine, Ananey-Obiri Daniel and Francis Poku, hereby affirm that this
document has been prepared and presented in agreement with the due procedure and academic
rules and conducts laid down by the School of Engineering as well as the Food Process
Engineering Department under the supervision of Dr. George Afrane. We also declare that, this
work is our own work, except where indicated by referencing.
STUDENT NAME: ACHARIBASAM VALENTINE
Signature: โฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆ..
Date: โฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆ..
STUDENT NAME: ANANEY-OBIRI DANIEL
Signature: โฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆ.
Date: โฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆ..
STUDENT NAME: POKU FRANCIS
Signature: โฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆ.
Date: โฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆ..
SUPERVISOR: DR. GEORGE AFRANE
Signature: โฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆ
Date: โฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆ..
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ACKNOWLEDGEMENT
Our profound gratitude goes first and foremost to God almighty who has given us the grace and
strength to accomplish this project successfully.
We wish to show our appreciation to our supervisor, Dr. George Afrane who has been of great
help to us and for his guidance.
We are also grateful to the whole teaching staff of the Food Process Engineering department for
their constructive criticisms which have helped shape and structure our activities to make this
project a success.
We also want to express our appreciation to our parents Mr. and Mrs Acharibasam Mathew, Mr.
Sampson Arthur Atta and Mr. Poku Francis for their financial and parental support during our
stay in Legon. God richly bless them.
We further extend our gratitude to all our colleagues, friends and most especially all our
course mates whose competiveness and encouragement have seen us through successfully.
Finally, we wish to express our gratitude to all those who contributed materially, spiritually
and morally to the success of this project.
To you all we say thank you and may the Almighty God richly bless you.
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EXECUTIVE SUMMARY
Sorbitol is a sugar alcohol produced by the reduction of the aldehyde group of glucose to a
hydroxyl group. It has a sweet and refreshing flavor and can be used as moisturiser, texturiser
and softener in the food industry. Other uses include vitamin C production and sorbose. Sorbitol
also has low calories and hence used as a substitute of sugar for diabetic patients.Pharmaceutical,
cosmetic and textile industries also use it as raw material for the production of other products
such as body cream and cough syrups.
The annual sorbitol importation of sorbitol into Ghana is reported to be increasing annually,
showing that the product is in increasing demand. The most basic raw material for sorbitol
production is starch, and starch is generally obtained from cereals and tubers. Cassava however
is known to be a source of low cost starch due to the less labour and resources required in its
production and processing. Cassava farmers in the country have been experiencing huge losses
due to lack of market for their produce. This project is intended to utilise the readily available
raw material to produce sorbitol to meet its increasing demand.
The plant is designed to produce 3,000 kg of sorbitol every day and will operate for 300 days
annually. An annual production of 900,000 kg of sorbitol crystals is expected to be produced.
The plantโs service life is estimated to be 20 years with an initial capital investment and working
capital of GHศป 4,754,314 and GHศป 838,997 respectively. The Net Present Value (NPV) of the
project from the sensitivity analysis is GHศป2,783,058 and a payback period of 1.9 years. The
Discounted Payback Period (DPB) is 3.6 years whereas the Rate of Return on Investment (ROI)
is 52.4%. A 50 kg box of sorbitol will be sold for GHศป 430. The above values indicate that, the
project is a profitable venture and worth executing.
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Table of Contents EXECUTIVE SUMMARY ......................................................................................................................... iii
CHAPTER ONE ......................................................................................................................................... 1
1.0 INTRODUCTION ........................................................................................................................... 1
1.1. DESCRIPTION OF KNOWN PROCESSES ............................................................................... 3
1.1.1. BIOLOGICAL PROCESS .................................................................................................... 3
1.1.2. ELECTROLYTIC METHOD ............................................................................................... 4
1.1.3. CATALYTIC HYDROGENATION METHOD .................................................................. 4
1.2 PROCESS SELECTION .................................................................................................................... 5
1.3 GENERAL BLOCK FLOW DIAGRAM FOR SORBITOL PRODUCTION ............................. 6
1.3.1. STARCH EXTRACTION .......................................................................................................... 7
1.3.1.1 PROCESS DESCRIPTION FOR STARCH EXTRACTION .................................................. 8
1.3.2 GLUCOSE SYRUP PRODUCTION ......................................................................................... 13
1.3.2.1 ALTERNATIVE PROCESS DESCRIPTION ........................................................................ 13
1.3.3 SORBITOL PRODUCTION .............................................................................................. 20
THE PROCESSING LINE (PRODUCTION OF SORBITOL) ......................................................... 20
CHAPTER TWO ...................................................................................................................................... 25
2.0 OBJECTIVES ..................................................................................................................................... 25
2.1 GENERAL OBJECTIVES ............................................................................................................... 25
2.1 SPECIFIC OBJECTIVES ........................................................................................................... 25
CHAPTER THREE .................................................................................................................................. 26
3.0 MASS AND ENERGY BALANCE ............................................................................................... 26
3.1 MASS BALANCE ............................................................................................................................ 26
3.1.1 STARCH PLANT MASS BALANCE ............................................................................... 28
3.1.2 GLUCOSE PLANT MASS BALANCE ............................................................................. 31
3.1.3 SORBITOL PLANT MASS BALANCE............................................................................ 33
3.2 ENERGY BALANCE ................................................................................................................ 35
3.2.1 GENERAL ASSUMPTIONS FOR ENERGY BALANCE CALCULATIONS .................... 35
CHAPTER FOUR ..................................................................................................................................... 37
4.0 BASIC DESIGN OF ALL PROCESS EQUIPMENTS .................................................................. 37
4.1.1 STARCH PLANT PROCESS FLOW DIAGRAM ................................................................... 37
4.1.2 GLUCOSE PLANT PROCESS FLOW DIAGRAM ................................................................ 38
4.1.3 SORBITOL PREPARATION PLANT PROCESS FLOW DIAGRAM ................................ 39
4.2 SPECIFICATIONS OF PIPELINES AND OTHER CONNECTORS ................................................. 40
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4.2.1 PLANT LAYOUT ...................................................................................................................... 40
4.2.2 SPECIFICATION OF EQUIPMENT PIPING .......................................................................... 46
4.2.3 SPECIFICATIONS OF THERMAL INSULATION ......................................................... 46
CHAPTER FIVE ........................................................................................................................................ 49
5.0 EQUIPMENT DESIGN ........................................................................................................................ 49
5.1 DESIGN OF TOP SUSPENDED MOTOR CYLINDRICAL-SCREEN CENTRIFUGE ......... 49
5.1.1 GENERAL CONSIDERATIONS IN DESIGNING TOP SUSPENDED MOTOR
CYLINDRICAL-SCREEN CENTRIFUGE ........................................................................................... 50
5.1.2 GENERAL FEATURES OF THE TOP SUSPENDED MOTOR CYLINDRICAL-
SCREEN CENTRIFUGE ................................................................................................................... 52
5.1.3 WORKING PRINCIPLE OF TOP SUSPENDED MOTOR CYLINDRICAL-SCREEN
CENTRIFUGE ........................................................................................................................................ 52
5.1.4 DESIGN CONCEPTS......................................................................................................... 54
5.1.4.1 COMPONENTS THE TOP SUSPENDED MOTOR CENTRIFUGE ............................... 54
5.1.4.2 INTERNAL VIEW OF THE TOP SUSPENDED CYLINDRICAL-SCREEN
CENTRIFUGE. ................................................................................................................................... 56
5.1.4.3 PARTS OF CENTRIFUGE ................................................................................................ 57
5.1.4.4 SANITARY DESIGN CONSIDERATIONS ......................................................................... 58
5.1.4.5 DESIGN MATERIALS ...................................................................................................... 58
5.1.4.6 DESIGN EQUATIONS AND DIMENSIONS OF THE TOP SUSPENDED MOTOR
BASKET CENTRIFUGE ................................................................................................................... 59
5.2 DESIGN OF ADSORPTION COLUMN ............................................................................................. 61
5.2.1 ADSORPTION MECHANISM ..................................................................................................... 61
5.2.1.1 ADSORPTION EQUILIBRIUM ............................................................................................ 62
5.2.2 BREAKTHROUGH CURVE ........................................................................................................ 63
5.2.3 BACKWASHING .......................................................................................................................... 64
5.2.4 COMPONENTS OF ADSORPTION COLUMN .......................................................................... 64
5.2.4.1 COLUMN (ADSORBER) ...................................................................................................... 64
5.2.4.2 ADSORBENT BED DESIGN ................................................................................................ 65
5.2.4.3 METALLIC SIEVE ................................................................................................................ 66
5.2.4.4 INFLUENT AND EFFLUENT PIPES ................................................................................... 66
5.2.3 ADSORPTION COLUMN DESIGN ......................................................................................... 67
5.3 DESIGN OF A HYDROGENATION REACTOR ...................................................................... 72
5.3.1 OPERATING PRINCIPLES ...................................................................................................... 72
5.3.2 MAIN COMPONENTS OF THE REACTOR .......................................................................... 73
5.3.3 MATERIALS FOR CONSTRUCTION .................................................................................... 73
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CHAPTER SIX ......................................................................................................................................... 76
6 PROCESS CONTROL ....................................................................................................................... 76
CHAPTER SEVEN .................................................................................................................................... 78
7.0 GENERAL CONSIDERATION ..................................................................................................... 78
7.1. SITE SELECTION .......................................................................................................................... 79
7.2 SANITATION DESIGN CONSIDERATIONS ........................................................................... 81
7.3 SAFETY CONSIDERATIONS IN THE PLANT ........................................................................ 81
7.4 GIPC LAWS IN RELATION TO INVESTMENT, CAPITAL AND LABOR EMPLOYMENT
............................................................................................................................................................ 81
7.5 UTILITIES .................................................................................................................................... 82
CHAPTER EIGHT ................................................................................................................................... 83
8.0 ECONOMIC ANALYSIS .................................................................................................................. 83
8.1 TOTAL CAPITAL INVESTMENT ........................................................................................... 84
8.2 ESTIMATION OF FIXED CAPITAL INVESTMENT .............................................................. 85
8.2.1 EQUIPMENT COST ............................................................................................................... 85
8.2.2 TOTAL IMPORTED EQUIPMENT COST ......................................................................... 87
8.2.3 COST OF LOCALLY ACQUIRED AND FABRICATED EQUIPMENT ........................ 90
8.2.4 TOTAL PURCHASED EQUIPMENT COST (TPE) ........................................................... 91
8.2.3 COST OF LAND ...................................................................................................................... 93
8.3.3 ESTIMATION OF FIXED CAPITAL INVESTMENT ....................................................... 93
8.3 ESTIMATION OF WORKING CAPITAL.................................................................................. 93
8.4 ESTIMATION OF TOTAL PRODUCTION COST (TPC) ....................................................... 95
8.4.1 ESTIMATION OF MANUFACTURING COST (๐๐) ........................................................ 95
8.4.2 DIRECT PRODUCTION COST ............................................................................................ 95
8.5 PROFITABILITY ANALYSIS ................................................................................................... 105
8.5.1 ESTIMATION OF ANNUAL REVENUE ........................................................................... 105
8.5.2 GROSS PROFIT (PG) ............................................................................................................ 108
8.5.3 TAXABLE INCOME (R) ...................................................................................................... 108
8.5.4 INCOME TAX (T) ................................................................................................................. 108
8.5.5 ANNUAL PROFIT AFTER TAX (P) .................................................................................. 108
8.5.6 ANNUAL CASH FLOW (CF) .............................................................................................. 109
8.5.7 NET PROFIT (PN) ................................................................................................................. 109
8.6 FINANCIAL APPRAISAL .......................................................................................................... 110
8.6.1 PAYBACK PERIOD (PBP) .................................................................................................. 110
8.6.2 DISCOUNTED PAYBACK PERIOD .................................................................................. 110
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8.6.3 CUMULATIVE CASH FLOW ............................................................................................. 111
8.7 SENSITIVITY ANALYSIS .......................................................................................................... 111
8.7.1 INTERNAL RATE OF RETURNS (IRR) ........................................................................... 112
CHAPTER NINE .................................................................................................................................... 114
9.0 CONCLUSION AND RECOMMENDATION .............................................................................. 114
REFERENCE .......................................................................................................................................... 115
APPENDIX A .......................................................................................................................................... 119
MATERIAL BALANCE ........................................................................................................................ 119
ENERGY BALANCE .............................................................................................................................. 149
ENERGY BALANCE FOR THE SORBITOL PREPARATION PLANT ....................................... 160
APPENDIX B ........................................................................................................................................... 174
PIPING AND FRICTION LOSSES CALCULATIONS .......................................................................... 174
PIPING AND FRICTION LOSSES CALCULATIONS (GLUCOSE PLANT) ...................................... 179
LENGTH OF PIPING IN THE SORBITOL PLANT ......................................................................... 185
APPENDIX C ........................................................................................................................................... 189
DETAILED EQUIPMENT DESIGN CALCULATION ................................................................. 189
TOP SUSPENDED MOTOR CYLINDRICAL-SCREEN CENTRIFUGE CALCULATIONS . 189
ADSORPTION COLUMN DESIGN CALCULATIONS ............................................................... 198
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CHAPTER ONE
1.0 INTRODUCTION
Sorbitol was first isolated from the juice of mountain ash berry (Srobus Americana, S. decora ) in
1872 by the French chemist, Joseph Bovssingavit. Schwartz & Whistler (2009) explained that
sorbitol like other sugar alcohols (polyol) such as xylitol, matitol is found naturally in some
plants, but commercial extraction is not feasible.
For industrial purposes it is done through hydrogenation of the glucose. Sorbitol is used as
sweetener, moisturizer, texturizer and softener in the food industry. It is also used for vitamin C
production and sorbose (Barros et al. 2006). Sorbitol is widely used as an additive in foods,
drugs, and cosmetics, and it is also used as an intermediate in ascorbic acid synthesis (Mishra et
al. 2012).
The most basic and important raw material for sorbitol production is starch. Starchy foods have
been utilized by humans for centuries. Traces of industrial starch production were found in
Patrician Torlonia Catoโs writing in 170 BC. The writing was used for starch separation from
grains by the Romans. The main sources of starch are barley, rice, wheat, sweet potatoes, cassava
and corn, with cassava starch gaining popularity in the starch industry. Cassava is widely valued
as a low cost carbohydrate source for urban consumers Hillocks (2002) and the food security that
it provides (Siritunga and Sayre, 2003).
The biological characteristics of cassava, its ability to survive after cultivation, and the viability
of its cuttings have contributed greatly to its spread (Lebot 2009). The carbohydrate values
present in cassava are consistent which range from 32% to 35% on fresh basis and 80% to 90%
on dry matter basis (Montagnac et al., 2009; Zvinavashe et al., 2011). Cassava contains about
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1-2% protein as compared to maize and sorghum which have about 10g protein/100 g fresh
weights (Charles et al., 2005; Montagnac, 2009). This makes it suitable for sorbitol production.
The crop ranks first among its root and tuber contemporaries inโterms of production in Ghana
(13,504 MT in 2010) and has an estimated per capita consumption of 152.9 kg/year (MOFA,
2010 and FAOSTAT, 2013).
Glucose (๐ถ6๐ป12๐6) also known as dextrose is a simple monosaccharide and it is either produced
by photosynthesis in plants or by hydrolysis of starch (acid or enzyme). The production of
glucose syrup with high Dextrose Equivalent (DE) value, which is the amount of soluble sugar in
the glucose is very important in the conversion to sorbitol. That is the higher the DE value of the
glucose, the greater amount of sorbitol produced.
The industrial processing of starch to sugars can be carried out either by acid or enzymatic
hydrolysis or the combination of both processes. However, the enzymatic hydrolysis is preferred
to the acid hydrolysis, since it produces high yields of desired products and less formation of
undesired products such as toxics compounds (Sanjust et al. 2004). The enzymatic hydrolysis
method is the one employed in this process.
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1.1. DESCRIPTION OF KNOWN PROCESSES
There are three known processes used industrially in the production of sorbitol. These processes
are namely;
i. Biological process of sorbitol production
ii. Synthesizing of sorbitol by electrolytic method and
iii. Catalytic hydrogenation of glucose
1.1.1. BIOLOGICAL PROCESS
The biological method of sorbitol production uses the bacteria Zymomonas mobilis to convert
glucose and fructose to sorbitol and gluconic acids. The bacterium achieves the result through
the reactions catalyzed by Glucose-Fructose Oxyreductase (GFOR); an enzyme with tightly
coupled NADP (Silveira et al. 1999; Wisbeck et al 1997). Sorbitol yield from the use of the
bacteria alone is relatively low due to the formation of gluconic acids and ethanol (Viikari 1984).
Thus to improve the sorbitol yield, various cell permeabilisation methods are evaluated by
releasing soluble cofactors necessary for the activation of the enzyme on the Enther Doudorulf
pathway. This resulted in then reduction of the ethanol produced and sorbitol was increased from
89- 100%.
DISADVANTAGES OF THIS METHOD
The Biological method can be classified as an environmentally friendly means of producing
sorbitol. However, the demerits are numerous.
i. There is a high cost involved in the cultivating and the growth of the bacteria.
ii. Relatively low yield of sorbitol
iii. Increased operational difficulties
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1.1.2. ELECTROLYTIC METHOD
According to Kassim and Rice (1980), increasing application of electrochemical synthesis has
raised attention to the synthesis of sorbitol by electrolyzing glucose. During the electrolysis,
glucose and Na2SO4 electrolyte are filled in the cathode part of the electrolytic cell. Potassium
hydroxide (KOH) was placed in the anodic chamber with a cation exchange membrane
separating the two medium. The electrolyte is circulated using pumps leading to the formation of
sorbitol.
Sorbitol production by electrolytic reduction of glucose initially used zinc (Zn) and lead (Pb)
cathodes. However, the lead and zinc electrodes corrode faster due to their soft properties and
low mechanical resistance. Moreover, both metals are poisonous and so the product cannot be
used in the food and pharmaceutical industries.
Presently, hydrogen-storage alloy is rapidly developed and widely used but there are no reports
on preparation of sorbitol using a hydrogen-storage alloy reduction electrode.
When optimal conditions of electrolysing glucose with alloy electrode are reached, the efficiency
of sorbitol conversion is a little over 90%. The optimal conditions are a current density of 8mA,
a voltage of 4.0 - 5.0 mV at a temperature of 30 C and a pH of 8.0.
1.1.3. CATALYTIC HYDROGENATION METHOD
Approximately 700,000 tonnes of Dโsorbitol are synthesized each year worldwide by the
catalytic hydrogenation of Dโglucose. This represents an inexpensive, abundant feedโ stock
obtainable from renewable resources such as starchโ containing crops or.
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1.2 PROCESS SELECTION
Catalytic hydrogenation of glucose has been selected for the process because of the high yield of
sorbitol which is about 100%, the lower maintenance cost and the fast conversion rate it gives
compared to the other forms of sorbitol production.
The production of sorbitol using the catalytic hydrogenation of glucose syrup involves three
successive steps;
i. Starch extraction
ii. Dextrose preparation
iii. Sorbitol production
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1.3 GENERAL BLOCK FLOW DIAGRAM FOR SORBITOL PRODUCTION
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1.3.1. STARCH EXTRACTION
The process selected is with the aim of producing quality starch suitable for sorbitol production
and to enhance a high yield of hygienic and safe starch. The most important quality standard of
cassava starch for sorbitol production is that, it should not contain no color, odor or any other
impurities undesirable to the subsequent process in sorbitol production.
The unit operations selected is based on the varied starch production processes available and
improved indigenous technologies to aid the production of pure cassava starch with minimal
impurities. The plant design will also take into consideration the sanitation of the plant, the cost
of production in terms of energy consumption and operate in a safer manner, providing an
acceptable hazard risk to the plant employees and the public.
STARCH EXTRACTION BLOCK FLOW DIAGRAM
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1.3.1.1 PROCESS DESCRIPTION FOR STARCH EXTRACTION
RAW MATERIAL (CASSAVA ROOTS) PREPARATION
The cassava roots are harvested and sold to factories within 24 hours to reduce the tendency of
getting low yields of starch and also reduce microbial load on the tubers .Mature roots can range
in starch content from as low as 15% to as high as 35%, depending on the climate and harvest
time. Starch content reaches a maximum at the end of the rainy season. Less mature roots will be
lower in starch content and higher in water, while overly mature roots will be lower in recov-
erable starch content and have a woody texture, making starch processing difficult (Breuninger et
al. 2009). Cassava roots for this process must be of high quality, in good state of health, without
signs of rot, and from a well matured (10 to 12 months), low moisture variety since these factors
have a direct impact on product recovery rate and starch quality of the effluent starch cake
(Dziedzoave et al., 2006)
Upon arrival at the factory, the root tubers are delivered into storage areas of cemented floors.
The tubers are then sampled and the sand load the tubers as wells as the starch content is
estimated. The starch content is estimated with a method adopted from the potato industry, using
a Rieman balance. The method of starch estimation is easy; 5Kg of clean root tubers is weighed
with Rieman balance then the same sample is weighed in water. The apparent densities of the
tubers in air and in water are both determined, and the correlation between the apparent density
and starch content are calculated (Breuninger et al. 2009). The sand load is determined by
weighing a sample of the dirt tubers, after which the sample is washed and weighed. The
difference between the weights tells the sand load on the tubers. This is to aid mass balancing.
The quality assurance department is responsible for doing these analyses.
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The starch content of the received cassava shows the quality of the cassava and can hence be
used to determine how much farmers are paid in other to encourage good farm practices and
quality cassava production.
ROOT SORTING
Sorting is the separation of foods into categories on the basis of a quantifiable physical property.
Sorting of foods is based on physical properties such as the size, color, weight and shape. In this
case, the sorting phase will include removing produce with surface deformities or blemishes,
dried or woody tubers and foreign / unwanted objects. Insect infested and bruised cassava tubers
are removed before they are delivered into the hopper.
ROOT WASHING
The roots are then feed into a root hopper and channeled into a sand removal drum which
consists of rotating inclined squirrel cage which is used to remove loose sand, sticks and other
unwanted materials. Roots after the screening stage are transported into a paddle washing
chamber where they are washed and moved along the process line by rotary washers. The paddle
washing machine combines flushing with a low pressure water and continuous removal of dirt
and peel. When the washing is done efficiently it reduces the burden on subsequent refining
processing. This makes the washing one of the most important stages of the process.
PEELING
The washing and peeling equipment are joined in a way that the tubers move immediately from
the washing at low water level to the peeling compartment. Thus a cylindrical rotating root
washing peeling drums. Water recycled from downstream processes is used during the peeling to
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wash the peels out and to aid the peeling process. The waste water can be treated for reuse since
it will have a high BOD and this poses environmental issues when discharged into drainage
systems. The peels contain high levels of protein, color and odors which are not desired in the
final product, hence the washing peeler must be efficient. The washed and peeled tubers are
conveyed on an inspection belt where humans do inspection before being sent to the pre-cutter.
In order to feed the raspers properly, the roots are chopped into pieces.
RASPING
Rasping is the first step in the starch extraction process. The goal is to open all the tuber cells, so
that all the starch granules trapped within the tubers are released for extraction. The rasping is
the most important part of the starch extraction process. After peeling and inspection, the roots
are chopped into pieces of 1-2 centimeters in size and fed into the curved mesh crusher. The
curved mesh crusher consists of series of rasping stages which gives it a high efficiency in
releasing the starch granules in the cassava completely. The machine adopts production works of
multi-crushing, multi-washing, multi-filter and multi-extrusion which ensures that minimum
quantities of starch remain in the cassava residue. And efficiently separate the starch slurry form
the cassava residue. The crushing is done by series of saw tooth rasping drums which has high
intense attrition and converts cassava into pulpy slurry. The pulpy slurry consists of fruit juice,
pulp and starch. Water recycled from subsequent stages is used mix the slurry for efficient
filtering. When the root tubers are rasped, the toxic hydrogen cyanide and cyanohydrin which
hinder the utilisation of cassava, are released and will be removed through subsequent processes
in the starch extraction and sorbitol production processes. The cyanogen in cassava are
hydrolyzed into volatile free cyanide by allowing contact between the cyanogenic substances
localised in the vacuoles of the cells with hydrolysing enzymes in the cell walls. This can be
11
achieved by damaging the cells mechanically or fermenting. The rasping stage is therefore very
essential.
EXTRACTION
The subsequent processes after rasping follow each other immediately to prevent fermentation.
The cell juice is rich in sugar and protein. When opening the cells, the juice is instantly exposed
to air and reacts with the oxygen, forming colored components, which may adhere to the starch.
Food grade sulphur dioxide gas or sodium-bisulphite-solution therefore has to be added. The
great reduction potential of the sulphur compounds prevents discoloration. Sufficient sulphur has
to be added to turn the juice and pulp light yellow (Sriroth et al. 2000).
The fresh rasped root slurry from the rasper is then pumped through a series of coarse and fine
extractors, where the starch granules are flushed out and the fiber is removed by screens arranged
conically in continuous centrifugal perforated baskets. Starch slurry exiting the coarse extractor
equipped with a filter cloth and a screen with an aperture of 150 microns (100 mesh) to 125
microns (120 mesh) still contains a large amount of fine fiber which must be removed in a fine
extractor equipped with a finer screen (140โ200 mesh) (Breuninger et al. 2009). Pulp from the
coarse extractor is repeatedly re-extracted to achieve minimal loss of starch trapped in the moist
pulp (60โ70% moisture content and 45โ55%, dry basis, starch content). This is done by
rechanneling the coarse extract to the rasper.
HYDROCYCLONE SEPARATION
Starch slurry during transportation from each extraction unit to a separator is passed through
hydrocyclones for complete sand removal. With hydrocyclones, it is feasible to reduce fiber and
juice to low levels with minimum fresh water used. Increasing the number of hydrocyclone
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refining steps may accomplish considerable savings of fresh water. This is one of the advantages
of using hydrocyclones. Hydrocyclones use centrifugal forces to classify fluid particles
according to their densities. Starch slurry particles have different densities which makes
hydrocyclone use feasible
Table 1.1 Starch slurry particle and density.
PARTICLE DENSITY g/ml
Starch 1.55
Water 1.00
Soil, sand Above 2
STARCH SEPARATION
Starch slurry received from fine extraction has a concentration of 10 to 17 ยฐBรฉ. Water is then
separated from the starch slurry, increasing the concentration to 18 to 20 ยฐBรฉ, using a separator
(Breuninger et al. 2009). The separator also uses centrifugal forces for the concentration of the
starch slurry.
DEWATERING
The semi-cake starch slurry is then pumped into a horizontal centrifuge for water to be removed.
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1.3.2 GLUCOSE SYRUP PRODUCTION
This section of the project is aimed at producing dextrose with high purity and high Dextrose
equivalent (DE) that is 98%. This is needed for the catalytic hydrogenation of the glucose to
produce sorbitol.
Cassava starch, like any other starch is a polymer of glucose. The glucose units are joined in a
chemical bond one carbon atom and carbon-4. Amylopectin and amylose are the two polymers
of starch. Aiyer et al.( 2005) explained that starch is found in nature as insoluble, non-dispersible
granules that can be hydrolysed. Liquefaction and saccharification are the main steps in the
hydrolysis of starch. The starch slurry is composed of 70% water (w/w).
The hydrolysis of starch can be achieved by three methods namely;
1. enzymatic hydrolysis
2. enzymatic- acid hydrolysis
3. acid hydrolysis
The resulting dextrose is purified to remove fats, protein and other impurities through ion
exchange after the adsorption process.
1.3.2.1 ALTERNATIVE PROCESS DESCRIPTION
ACID HYDROLYSIS
Firstly, the hydrolysis of starch was achieved by boiling raw starch in sulphuric acid to give
sweet syrup. The commonly used acid is hydrochloric acid which is done at temperatures of 130-
170ยฐC with subsequent partial neutralization. The acid serves as the thinning agent. In this
process, the starch slurry is acidified with the hydrochloric acid and pumped through steam-
heated pipes where the conversion takes place, which is from starch to glucose syrup. The use of
14
this method has some advantages and disadvantages. Fontana et al.( 2008) explained that, the use
of hydrochloric acid produces toxic fumes. Secondly, after neutralisation, it becomes necessary
to remove the undesirable irons, salt with high cost iron exchange resin. The process is also
known to give low yield of glucose accompanied by food safety issues.
ACID-ENZYME HYDROLYSIS
The acid-enzyme hydrolysis process involves the use of both acid and enzyme. The glucose
syrup is manufactured by hydrolysis of the starch with acid and completing the hydrolysis by
using one or more enzymes. The enzyme is added after the starch has been cooked and cooled to
100 โ 95ยฐC (Aiyer et al. 2005). This process also does not yield glucose syrup with high
Dextrose Equivalent (DE).
ENZYMATIC HYDROLYSIS OF STARCH
Enzymatic hydrolysis of starch is the method employed in the glucose production process in this
project because this method yields high DE and it poses no safety issues. Further details are
provided below.
15
GLUCOSE PRODUCTION BLOCK FLOWW DIAGRAM
16
LIQUEFACTION
The process involved in the conversion of starch to dextrose begins with liquefaction of the
starch slurry at a temperature of 100- 105oC at 1at m pressure using steam as a source of energy.
The liquefaction process is a thinning process and also this process produces small amount of
glucose as well. Gelatinisation, the initial process in the liquefaction process leads to the
absorption and swelling in the presence of water and heat.
During the liquefaction process , the hydrogen bond between the starch mixture weakens and this
permits them to swell and gelatinized (Aiyer et al. 2005). The enzyme, ฮฑ-amylase is added from
the alpha amylase dosing tank. The pH of the gel is adjusted to about 6-6.5. The pH of the
mixture is not allowed to fall the below otherwise the amylase will be denatured. The
temperature is kept at 90oC for the optimum performance of the enzyme, ฮฑ-amylase. Calcium
ions, about 50ppm quantities are added to stabilise the enzyme in the Continuous Stirring
Reaction Tank (CSRT). The enzyme hydrolyses the ฮฑ-1, 4- glucosidic bonds of amylose,
amylopectin, glycogen to produce oligosaccharides and small amount of glucose. This process
reduces the viscosity of the mixture and makes it more liquid. The starch is broken down both by
the shearing force of the paddles of the propeller and by the action of the enzyme. After the
initial liquefaction for about 10 minute, the mixture is cooled to 97oC and transferred to a vessel
where the solution is held for 90 minutes to reach Dextrose Equivalent of about 10-12.
SACCHARIFICATION
The liquor is then pumped to the saccharification reactor which also has a continuous stirring
paddle. The glucoamylase enzyme is added from a dosing tank after the temperature has been
17
adjusted to 55-60oC and maintained at that temperature for the rest of the dwelling time. The pH
also is reduced to about 3.5-4.5 by the addition of sulphuric acid from an upward connected pipe.
These conditions are necessary in the inactivation of the ฮฑ-amylase and for optimal performance
of the enzyme glucoamylase. The glucoamylase is added to hydrolyse both the ฮฑ-1,4 linkage and
the ฮฑ-1,6 bonds to completely hydrolyse the dextrin (Aiyer et al. 2005). According to the amount
of enzyme added this way, the saccharification time ranges from 24 to 96 hours when a
maximum conversion of 95-96 DE is attained. The resulting syrup having DE value between 96
and 97 can be used to produce crystalline. The liquefied mixture flows via an expansion valve
which maintains the pressure in the system to the acidification tank,
The equation below describes the saccharification process
(๐ถ6๐ป10๐5)๐ + ๐๐ป2๐ ๐๐๐ข๐๐๐๐๐ฆ๐๐๐ ๐ ๐๐ถ6๐ป12๐6
PURIFICATION
The purification process involves three processes to obtain dextrose of high purity of about 98%
from the crude syrup. First, the removal of the protein content which is about 1.2% in the starch
slurry by using the isoelectric precipitation. This is followed by the adsorption which involves
the use steam activated carbon to treat the glucose syrup to remove colorants etc. The de-
ionisation of the dextrose syrup using ion exchange resin follows thereafter.
ACIDIFICATION
The acidification involves the use of dilute food grade sulphuric acid to reduce the pH of the
glucose syrup to about 3-3.3. The acidification process uses the isoelectric precipitation of
protein. The reduction in pH causes the protein content in the glucose syrup to coagulate and
precipitate and this is subsequently removed during the adsorption and ion exchange process.
18
ACTIVATED CARBON TREATMENT
The syrup is passed through a fixed bed of activated carbon for clarification and bleaching. The
temperature in the carbon column is maintained at 150โ170ยฐF (69โ77ยฐC) with a typical contact
time of 90โ120 minutes for optimum removal of impurities. The activated carbon is made of
Granulated Activated Carbon (GAC) packed in columns. The column removes precursor colours
and flavours as the syrup flows through the packed column accompanied by the accumulation of
those substances at the surface of adsorbent phase. The spent carbon is removed, regenerated in a
furnace and repacked at the top of the column. The column also adsorbs some coagulated and
precipitated protein from the syrup.
ION EXCHANGE
The de-ionisation process is crucial in removing calcium ions and other ions to prevent poisoning
of the catalyst used in the hydrogenation process. The process also improves the colour and
stability of the dextrose syrup by removing components that could otherwise participate in a
Maillard reaction with the reducing sugars(Hobbs 2009). Ion exchange resins are synthetic
organic polymers containing functional groups that exchange mobile ions in a reversible reaction
based on affinities. Synthetic polystyrene with sulphonate groups to form cation exchangers or
amine groups to form anion exchangers are used. Exhaustion of demineraliser is usually detected
by an electrical conductivity cell installed at the outlet. When the conductivity rises to indicate
ionic break through, a regeneration cycle can be initiated automatically. Exhausted resin depleted
of desired ions and is regeneration by contact with a down flow of high concentration of the
desired ions.
19
ION EXCHANGE (CATION)
Spargers at the top of the fixed bed cation exchange column distribute the dextrose containing
undesired ions. The resin beds consist of styrene and divinylbenzene which have been activated
with sulphuric acid. During service, cations in the glucose are taken up by the resin while
hydrogen ions are released. This increased in hydrogen ions further helps in precipitating protein.
ION EXCHANGE (ANION)
The second bed consists of anion resin bed. Here, the anions are exchanged for hydroxide ions,
which react with the hydrogen ions to form water. The resins will have 0.1mm diameter. As the
solution pass down the resin bed it flows through the cross-linked polymer, bringing it into
intimate contact with the exchange sites.
20
1.3.3 SORBITOL PRODUCTION
The sorbitol plant is made up of three main sub sections. The preliminary section that deals with
the reception, storage of the glucose and removal of air from the syrup. The intermediary stage is
the stage where the main process of converting glucose syrup to sorbitol occurs and the post
processing stage where the produced sorbitol crystals are stored in warehouse.
THE PROCESSING LINE (PRODUCTION OF SORBITOL)
Preliminary
Reception of dextrose syrup
Intermediate
Hydrogenation
Evaporation
Crystallization
Filtration
Drying
Milling
Post Processing
Packaging
Storage
21
SORBITOL PROCESSING BLOCK FLOW DIAGRAM
22
RECEPTION OF GLUCOSE SYRUP
Dextrose syrup from the other processing line will be transported through pipes into the storage
tank. The tank will be elevated above the ground to make use of the effect of gravity when
discharging into the hydrogenation reactor. The tank will be stirred continuously whilst pumping
nitrogen gas into it to cause the escape of all air within the syrup. This is necessary to prevent
any explosion when the glucose syrup comes into contact with the hydrogen gas.
HYDROGENATION
Dextrose syrup is transported through pipes to the reaction chamber of the hydrogenation reactor.
The unsaturated compound (dextrose syrup) undergoes a reduction reaction in the presence of
hydrogen gas and a catalyst. Thus C6H12O6 reacts with H2 to produce C6H14O6. The reaction is
highly exothermic.
EVAPORATION
After hydrogenation when the dextrose has been converted to sorbitol and water, they are
pumped into storage tanks before being transported to the evaporator for excess water to be
removed. This is achieved by adding latent heat of vaporization to the solution. The removal of
the water results in a concentrated sorbitol. In the industry, numerous evaporators are used based
on the concentration of the liquid, temperature sensitivity of the material and the temperature and
pressure. The vacuum evaporator will be used because higher evaporation temperatures will
cause the caramelisation of the solution.
CRYSTALLIZATION
Concentrated sorbitol solution is transported to the crystallizer where solid particles are formed.
This is achieved by heating the solution to a higher temperature and cooling it. Crystals of
23
sorbitol in the mother liquor are the result of the process. This unit operation is important
because most of the chemical products are marketed in this form.
FILTRATION
The suspended sorbitol in liquid medium is transported for the separation of the crystals from the
mother liquor using a porous membrane that retains the sorbitol crystals and removes the filtrate.
The principle of constant pressure filtration will be employed.
DRYING
After filtering to get the sorbitol crystals, they are dried to reduce their moisture content to about
1%. This is the ideal moisture content of sorbitol sold on the market. Drying is achieved by
blowing hot air on the surface of the crystals till the desired moisture content is achieved.
MILLING
After drying, the crystals are transported on conveyors to the mill where they are ground by
passing them through a three pair of rollers. The fine ground crystals are graded and sent for
packaging.
PACKAGING
The desired weight of product will be weighed by the packaging equipment, filled into a food
grade polyethylene bag and sealed. The sealed product will be conveyed to the other side of the
packaging area where the sealed product will be placed in a fitting food grade paper cardboard.
The choice for selecting these packaging materials is below.
Food Grade Paper Packaging systems
It can be easily recycled.
24
Assembling and disassembling is fast and easy.
It is biodegradable
Takes less space in the warehouse.
It is cheaper compared to other packaging materials.
Arrangement and transportation are easier.
Food grade Polyethylene
Relatively cheaper.
Films are soft and clear.
Lowest softening and melt point (good for heat-sealing)
Fair moisture barrier.
STORAGE
The product will be stored on pallets in a ware house at ambient temperature.
25
CHAPTER TWO
2.0 OBJECTIVES
2.1 GENERAL OBJECTIVES
The objective of this project is to design a plant for the production of sorbitol from cassava. The
plant will operate continuously for 300 days in a year. It will receive 15 tonnes of cassava daily
to produce 6049 kg of starch slurry. This is converted to 5023 kg of glucose syrup. A final
product of 3,000 kg of sorbitol is obtained.
To realise this, the following objectives must be accomplished;
1) To design a plant to extract starch from cassava
2) To design a plant to convert cassava starch to glucose syrup
3) To design a plant for the conversion of glucose syrup to sorbitol
This project will require a Total Production Cost (TPC) of GHC 4,705,459 and a total capital
investment of GHC 5, 593,311. The plant has a payback period of almost two years, a discounted
payback period of 3.6 years, a rate of return on all investment of 52.4% and a profitability index
of 0.5. This indicates that the project is viable.
2.1 SPECIFIC OBJECTIVES
1) To provide a detailed process design, mechanical engineering design and a 3D design of
a sorbitol filtration equipment for filtering sorbitol crystals from the mother liquor.
2) To design in detail, an adsorption column for the purification of glucose syrup.
3) To design a hydrogenation reactor that will optimize energy consumption whilst
efficiently accelerating the conversion of dextrose to sorbitol.
26
CHAPTER THREE
3.0 MASS AND ENERGY BALANCE
3.1 MASS BALANCE
Material balance calculations are employed in tracing the inflow and outflow of material in a
process and thus establish quantities of components or the whole process stream. The production
rate is small because only about 7% of the sorbitol market in Ghana has been targeted. There will
be an increase as the product gains popularity on the market.
Table 3.1:Yearly importation of sorbitol into Ghana.
Source: United Nations Commodity Trade Statistics Database
Year Trade Value Weight (kg) Quantity
2005 $155,691 195,303 195,303
2006 $179,988 309,624 309,624
2007 $271,188 409,334 409,334
2008 $448,470 579,213 579,213
2009 $389,264 633,599 633,599
2010 $520,772 757,294 757,294
2011 $811,341 1,017,120 1,017,120
27
The Global sorbitol demand was 1,829.6 kilo tons in 2013 and is expected to reach 2,337.2 kilo
tons by 2020 growing at a rate of 3.6% (European Association of Polyol Producers). There is an
increasing demand for sorbitol in Ghana, in 2011, a total of 1,017,120 Kg of sorbitol was
imported into Ghana and this figure is expected to increase (United Nations Commodity Trade
Statistics Database). Extrapolating using the values from the above table, 6,665,076kg of sorbitol
is expected to be imported into the country in 2017.
This plant is expected to produce 900,000kg of sorbitol annually, which is about 7 % of the
sorbitol expected to be imported into the country in 2017.
BASIS FOR MATERIAL BALANCE
(Total mass in) - (Total mass out) = (Total mass accumulated)
The processes are at steady states hence Accumulation= 0
(Total mass in) = (Total mass out);
(Component mass in) - (Component mass out) = (Component mass accumulated)
Assuming no accumulation,
(Component mass in) = (Component mass out)
The material balance calculations are clearly illustrated in Appendix A
28
3.1.1 STARCH PLANT MASS BALANCE
Assumptions for the starch factory;
The plant capacity is 15 tonnes of cassava per day
The peel part of the root comprises about 15% of the total weight of the root (Alves,
2002).
The starch content of cassava ranges from 32% to 35% on fresh basis, and when the tuber
is passed through the plant 26% of the starch present in the tubers is expected to be
extracted
The final pulp cake is diluted to 70% water content. That is 6049 kg of starch slurry.
SUMMARY OF MASS BALACE (STARCH PLANT)
29
2ND EXTRACTIONH
PUPL PRESSI
COARSE STARCH Hi111330.29Kg/day
WATER, Hi245% OF Hi1
5098.63 Kg/day
PULP Ho29064.23 Kg/day
STARCH SLURRY, Io131% OF Ho2
2809.91 Kg/day
FINE SLURRY Ho165% OF Hi1
7364.69Kg/day
PULP CAKE Io26254.32 Kg/day
30
31
3.1.2 GLUCOSE PLANT MASS BALANCE
ASSUMPTIONS
The plant receives 6049Kg of starch slurry per day to produce 5023 kg of glucose syrup.
The starch slurry contains 70% water and the final glucose syrup is 50% water
SUMMARISED MASS BALANCE (GLUCOSE PLANT).
32
33
3.1.3 SORBITOL PLANT MASS BALANCE
34
Summary of the total material balance of the sorbitol preparation process
35
3.2 ENERGY BALANCE
The energy balance was calculated at the boundaries of a unit operations based on the law of
conservation of energy which states that energy can neither be created or destroyed but can be
transformed from one form to another.
3.2.1 GENERAL ASSUMPTIONS FOR ENERGY BALANCE CALCULATIONS
Steady State operation
Energy in = energy out + accumulation
Since there is no accumulation,
Energy in= energy out
โ๐ธ = ๏ฟฝ๏ฟฝ(โ๐ป + โ๐พ๐ธ + โ๐๐ธ) ยฑ ๐ ยฑ ๐
At steady state, โ๐ธ = 0
There is no movement of equipment, velocity is zero hence โ๐พ๐ธ = 0
โ๐๐ธ = 0
The above equation reduces to
๐ = ๏ฟฝ๏ฟฝโ๐ป + ๐
For the preliminary design calculations and equipment sizing, the main forms of energy
considered is heat energy and mechanical energy, hence heat balance for various unit operations
involving heating are calculated. The mechanical energy requirement for various moving parts of
the process equipment were also determined. The table below shows the summary of the energy
used in the sorbitol production process.
36
Table 3.1 Summary of Energy Balance
PROCESS/EQUIPMENT ENERGY REQUIRED (MW)
Starch plant energy requirements
Total installed capacity 18
Glucose Production Plant Energy Requirements
Total energy required 5.2
Sorbitol Preparation Plant Energy Requirements
Total energy required 6.18
37
CHAPTER FOUR
4.0 BASIC DESIGN OF ALL PROCESS EQUIPMENTS
4.1.1 STARCH PLANT PROCESS FLOW DIAGRAM
38
4.1.2 GLUCOSE PLANT PROCESS FLOW DIAGRAM
39
4.1.3 SORBITOL PREPARATION PLANT PROCESS FLOW DIAGRAM
LEGEND
1. STORAGE TANK
2. 2.REACTOR
3. SORBITOL STORAGE TANK
4. EVAPORATOR
5. CRYSTALLISER
6. FILTRATOR
7. DRYER
8. ROLLER MILL
40
4.2 SPECIFICATIONS OF PIPELINES AND OTHER CONNECTORS
The design process of a new plant requires the engineers to focus on the arrangement of the
equipment and resources in a manner that will maximize efficiency, reduce cost of production,
and enhance the safety of both employees and the environment without compromising quality,
safety and acceptability of the end product. The site and equipment selection is therefore done
with the aforementioned factors in mind.
4.2.1 PLANT LAYOUT
The effectiveness of production in a food processing plant is dependent on the arrangement of
equipment and piping. The plant layout of processing equipment should be based on the
requirements of the flow of material, access to equipment, hygienic operations, process control
and maintenance (Maroulis and Saravacos 2003). The aims of a Plant layout design are primarily
to minimize unit cost, optimize quality, provide safe and convenient working environment,
control project cost, achieve production deadlines and promote effective use of operating
personnel, equipment, space and energy.
41
GENERAL PLANT LAYOUT
42
Table 4.1 Dimensions of proposed individual facilities and total areas covered
within plant
FACILITY DIMENSIONS (FT ร
FT)
TOTAL AREA
(FT2)
SECURITY CHECKPOINT (1,2 & 3) 5ร5 EACH 75
RAW MATERIAL RECEPTION 10ร15 150
CHANGING ROOM 25ร30 750
WASHROOM 19 ร25 475
CANTEEN 19ร32 608
EXPANSION AREA 100ร200 20,000
VISITORS CAR PARK 18ร25 450
WORKERS CAR PARK 19ร25 475
LOADING TRUCKS PARK (1 & 2) 19ร25 475
RECEPTION 8ร10 80
CLINIC 19ร22 418
WAREHOUSE 1 & 2 70ร100 7,000
ADMINISTRATION BLOCK 27ร38 1,026
QUALITY ASSURANCE DEPARTMENT 25ร30 750
ENGINEERING DEPARTMENT 19ร25 475
UTILITIES 19ร25 475
STARCH EXTRACTION AREA 66ร115 7,590
DEXTROSE PREPARATION PLANT 60ร100 6,000
SORBITOL PRODUCTION AREA 60ร100 6,000
WASTE TREATMENT PLANT 19ร22 418
ROAD WIDTH 15 wide 15
FACILITY SPACING 10 wide 10
TOTAL 53,715
43
STARCH PLANT LAYOUT
44
GLUCOSE PLANT LAYOUT
LV
D-2
CC
SATD-1
ADATSV
ACControl Room
Electical Room
100 ft
60
ft
LEGEND
LV โ Liquefaction vessel
SV- Saccharification vessel
D-1 Alpha Amylase Dosing Tank
D-2 Glucoamylase Dosing Tank
AT- Acidification Vessel
SAT- Sulphuric Acid Tank
AD- Adsorption Column
CC- Ion exchange Column (Cation)
AC- Ion Exchange Column (Anion)
45
SORBITOL PLANT LAYOUT
A= Hydrogen tankB=glucose storage tankC= heat exchangerD= hydrogenation reactor
E=Vacuum evaporatorF= crystallizer G= filtration unitH= Dryer
I= Roller millJ= Packaging equipment K= conveyor beltsL= pumps
A
10m
B
C
D
E FG
H I J
10m40m40m
20m
STORAGE ROOMS
HYDROGENATION ROOM
CRYSTAL FORMATION AREA
POST PROCESSING AREA GENERAL OFFICE
CONTROL ROOM
ENGINEERING ROOM
CHANGING ROOMS
KK
K
K
46
4.2.2 SPECIFICATION OF EQUIPMENT PIPING
All pipelines and connectors in this plant shall be made of stainless steel specifically the (ASTM
A270). This is because of its corrosive resistance and durability. Piping shall be designed in
accordance with the American Society of Mechanical Engineers (AMSE) standards and shall be
permanently identified by means of painted bands and letters.
i. All piping attached to equipment with rotating parts shall be adequately supported to
prevent excessive load on the equipment.
ii. In case stainless steel component is bolted between carbon steel flanges, stainless steel
gaskets shall be used.
iii. Gas piping shall be designed for good line drainage.
4.2.3 SPECIFICATIONS OF THERMAL INSULATION
i. To prevent heat loss in the plant, all pipes shall be insulated with resin bonded long fibred
mineral wool with a minimum weight of 100kg/m3.
ii. All indoor pipe works shall be finished with a 0.6mm plain aluminum jacketing whilst
outdoor pipes shall be 0.9mm plain aluminum jacketing.
47
PIPELINES SPECIFICATIONS AT PRODUCTION PLANT
Detail calculations for pipelines specifications are clearly illustrated in Appendix are clearly
illustrated in Appendix B
Table 5.2 Pipe length and pressure drop in the sorbitol processing section
FROM TO Schedule
number
length
(meters)
Internal
diameter
(mm)
pressure
drop (MPa)
Rasper Mixing tank 10 2 82.9 5.772
ร 10โ5
Mixing tank Extractor 10 5 82.9 1.35 ร 10โ4
Extractor Hydrocyclone
and separator
160 4 66.7 8.347
ร 10โ5
Separator Dewatering
machine
10 6 82.9 1.74 ร 10โ4
Mixing tank Liquefaction
vessel
5 8 84.7 1.98 ร 10โ4
Liquefaction
vessel
Saccharification
vessel
5 12 45 3.92 ร 10โ3
Saccharification
vessel
Acidification
tank
10 10 54.7 2.28 ร 10โ3
48
Acidification
tank
Adsorption
column
5 15 30.1 0.011
ร 10โ3
Adsorption
column
Cation exchange
column
10 15 54.7 3.39 ร 10โ3
Cation
exchange
Anion exchange
column
5 8 45 2.6ร 10โ3
49
CHAPTER FIVE
5.0 EQUIPMENT DESIGN
Engineering in food processing usually involve the manipulation of mass, energy and flow
processes to optimize the food manufacturing process either through biological, chemical,
mechanical or other approved scientific means to achieve desired results. This chapter is to
design equipment which use mechanical means of improving the sorbitol production process
5.1 DESIGN OF TOP SUSPENDED MOTOR CYLINDRICAL-SCREEN
CENTRIFUGE
In the sorbitol production process, the separation of sorbitol crystals from the magma after the
crystallization process is very essential. The separation can be done with a filter press or a
centrifuge. In our situation we are using a centrifuge.
The term centrifugal literally means moving away from the center, and a centrifuge is a machine
or equipment that uses centrifugal force for separating substances. This settling rate of particles
in fluids can be greatly influenced by the action of centrifugal forces as compared to
gravitational forces. The separation of the particles is achieved by accelerating the particles
mechanically in a circular force field. This power of the for exerted on the particles is
represented by equation 6.1, where r is the radius of rotation, m is the mass of particle, ๐ is the
angular velocity and Fc is the centrifugal force generated.
๐น๐ = ๐๐๐2 โฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆโฆ . .6.1
50
TYPES OF CENTRIFUGE
Over the years, there have been different designs of centrifuges applied in the food industry. The
two main types are sedimentation and filtration. A sedimenting centrifuge consists of a solid wall
cylinder or cone rotating about a horizontal or vertical axis, while the filtering centrifuge
contains a basket wall that is perforated and lined with a filter medium such as a cloth or a fine
screen. In this case the focus of this work is the filtering centrifuge. According to ( Afrane,
2012), the most common types of centrifuges are;
Tubular-bowl centrifuge
Disc-bowl centrifuge
Conical-screen centrifuge
Cylindrical-screen centrifugal filter
Ultracentrifuge
Gas centrifuge
The cylindrical-screen centrifuge filter is used in the sugar production industry to filter sugar
crystals from the massecuite. Sorbitol crystal filtration is similar to the sugar crystal filtration. So
the cylindrical-screen centrifuge is modified to function as a sorbitol crystals filter in this work.
5.1.1 GENERAL CONSIDERATIONS IN DESIGNING TOP SUSPENDED MOTOR
CYLINDRICAL-SCREEN CENTRIFUGE
In the designing of a centrifuge, there are general requirements of safety, good process
performance, low basket cost, high process and energy efficiency, with the overriding
requirement being safety. For centrifuges used within the European Economic Area, the
51
centrifuge must be designed to satisfy the requirements laid down in the type C standard
EN12547-2014, which specifies the safety requirements of common centrifuge. The standards
cover many aspects of a centrifuge design, however large aspects are about safety requirements,
protective measures and verification of mechanical hazards associated with ejection of part of the
rotating centrifuge basket. The materials and design selected must therefore satisfy these
requirements.
The filtration of the sorbitol crystals from the mother liquor from the crystallization unit is one of
the final processes in the sorbitol production process. The separation centrifuge therefore has to
be efficient in separating the crystals from the liquor.
The general considerations in designing the basket centrifuge are;
1) Safety and ease with which it is operated
2) Conformance to standards
3) Low human intervention
4) Efficient use of energy and power through reduction of friction to the minimum level
possible
5) Easily maintained and clean
6) Provision of large filtration surface for filtration.
7) Adoption to automated repetitive operations.
52
5.1.2 GENERAL FEATURES OF THE TOP SUSPENDED MOTOR CYLINDRICAL-
SCREEN CENTRIFUGE
The main features of the top suspended motor cylindrical-screen centrifuge are: a bowl or rotor
in which the mother liquor mixed with the sorbitol crystals to be separated is accelerated to
generate the centrifugal force, a feed inlet for introducing the material, a drive shaft, drive shaft
bearings, an electric motor to rotate the shaft and rotor, a casing or cover to isolate the separated
products and a frame to support and align these elements and casing-to-shaft seals.
5.1.3 WORKING PRINCIPLE OF TOP SUSPENDED MOTOR CYLINDRICAL-SCREEN
CENTRIFUGE
The top-suspended Motor Centrifuge is composed of rotary drum, motor, bearing seat, upper
housing, out housing, lower shaft, scraper group, feed pipe, wash pipe, cake discharge vent and
the liquid discharge. Its structure adopts cop located transmission system, vertical motor provides
the driving of drum through the coupling directly and the drum is fixed on the lower end of a
shaft.
The working principle of the top-suspended motor centrifuge: The motor on the top of the
equipment provides the driving force for producing the centrifugal forces. The motor drives the
rotary drum to turn. When the drum reaches feeding speed, the suspension (sorbitol magma) to
be separated will enter into the drum on a high speed from the feeding pipe. Feeding will stop
when the preset volume of the basket is reached. Then the speed of the drum will be raised to
generate enough centrifugal force for separation. Under the centrifugal force, the magma which
is a mixture of sorbitol crystals and liquor will be filtered by a filter cloth (filter screen). The
liquid phase will be thrown through the rotary drum hole to the empty chamber and discharged
53
through the liquid discharging pipe. The solid phase will be retained on the drum and form
cylindrical filter cakes which will then be washed by spraying water from the wash pipes. After
the washing, the filtration continues until the required moisture content is met, the speed will be
lowered for discharging. The centrifuge adopts dynamic or regenerative braking to achieve
obvious efficiency effects. When the required separation specifications are met, the breaking
system slows the rotating drum quickly to aid in the discharge process. The scraper device then
moves downward along the axial direction to scrape the filter cakes off the internal wall of the
drum and discharge them from the discharging pert on the bottom of the centrifuge. The
discharge filter cake is then taken to the next stage of the process.
The centrifuge adopts frequency motor driving and full-automatic repetitive operation. During
the automatic repetitive operation, the optimized operation is realized through frequency stepless
speed control so that the requirements of low speed feeding, high speed separating and low speed
discharging are met. The control is can be done by programmable logic controllers PLCs
systems. Time setting can be carried out for all working procedures and the real-time working
status of each procedure can be displayed on the operation screen.
54
5.1.4 DESIGN CONCEPTS
Fig 6.1 3D view of top suspended motor cylindrical-screen centrifuge
5.1.4.1 COMPONENTS THE TOP SUSPENDED MOTOR CENTRIFUGE
The various components of the top suspended motor basket centrifuge and their descriptions are
shown below;
Driving motor: this is a vertically placed motor which is connected to a shaft to drive the
filtering bowl.
Driving shaft: This is an elongated shaft that connects to the motor at the upper part and
to the rotary bowl/drum at the bottom. It has a plate-like structure along the middle part
which aids in the dispersion of the feed during filing.
55
Feed pipe: This is a pipe through which the feed is introduced into the centrifuge. The
feed is slanted at the lower end to direct the feed onto the plate-like collar on the shaft to
aid in the uniform distribution.
Bowl: the bowl is the filtering screen through which the liquid passes out and leaves the
cake inside the bowl. The bowl is connected to the driving shaft at the bottom and rotates
with the shaft during the filtration process.
Wash pipe: the wash pipe stretches from the upper housing downwards into the bowl. It
has orifices along its length through which the wash liquid is sprayed through onto the
filter cake.
Scraper group: this device is located at the upper housing and stretches downward to the
bowl. It has curve tip which is used to scrape off the cake from the filter screen. The
scraper group consists of a scraper tip as well as a hydraulic system which pushes the
scraper downward into the bowl to scrape off the cake and moves back up out of the bowl
when cake discharge is completed.
Outer housing: this is a cover that collects the filtrate and discharges it through the
liquid discharge. It is connected to the main frame of the centrifuge but not to the rotating
shaft and the bowl. It is the static part of the equipment.
Bearing seat: the bearing seat support the motor load while reducing friction. Bearing
guide the driving shaft
56
5.1.4.2 INTERNAL VIEW OF THE TOP SUSPENDED CYLINDRICAL-SCREEN
CENTRIFUGE.
Fig 6.2 Cross section view of centrifuge
57
5.1.4.3 PARTS OF CENTRIFUGE
Cover with feed pipe, wash pipe and scraper Housing of basket
Filter basket Shaft with filter basket
58
5.1.4.4 SANITARY DESIGN CONSIDERATIONS
Some of the sanitary equipment design considered includes:
Equipment is safe, easily cleaned and repaired
There should be no wooden parts
All surfaces of the machine are readily accessible and designed for quick dismantling
and reassembling with no tools, or with very simple tools
All metal surfaces in contact with the food product are made of food grade stainless steel
(316L).
Surfaces in contact with food are smooth, continuous, and without rot cracks or crevices
5.1.4.5 DESIGN MATERIALS
PARTS OF EQUIPMENT MATERIAL OF CONSTRUCTION
Driving Shaft Stainless steel
Motor Cast Iron
Housing Stainless steel
Filter Basket Stainless steel
Upper Frame Mild steel
Wash Pipe Copper
Feed Pipe Stainless steel
Scraper Stainless steel with plastic end
Support Table Mild steel
59
5.1.4.6 DESIGN EQUATIONS AND DIMENSIONS OF THE TOP SUSPENDED MOTOR
BASKET CENTRIFUGE
In the design of a basket centrifuge, the final design must balance the aforementioned competing
requirements, with safety requirements override all the other requirements. The material for the
centrifuge must be able to withstand the large stresses imposed on it by the centrifugal forces
generated by the rotary movement of the centrifuge basket.
A good understanding of the loads encountered by the basket during operation is a necessary first
step in the design process. The basket is subjected to a number of stresses during operation and a
full investigation of possible loads is necessary (Grimwood 2015).
i. Stresses due to centrifugal forces acting on the weight of the basket during rotation
ii. Stresses due to the centrifugal forces acting on the mass of the material placed in the
centrifuge to be separated.
iii. Stresses due to centrifugal forces acting on any unbalances.
60
MAIN PARAMETERS CHOSEN
Basket inside diameter, D (mm) 1200
Basket height, H (mm) 1000
Maximum speed, N (r/min) 1200
Separation factor
Maximum capacity of basket (Kg) 1000
Weight (kg) 2110
Capacity factor, CF (๐๐) 8934.39
Theoretical capacity of centrifuge ๐ธ๐ (๐๐/๐) 70581.7
Stress in the internal Basket ๐น๐(๐๐ต) 10606
Allowable thickness of material for
construction of internal basket ๐๐(๐)
0.06
Power requirement P (kW) 1507389.2
Twisting moment ๐ด๐(๐๐ต) 719725
Minimum diameter of shaft (๐) 6. ร ๐๐โ๐
61
5.2 DESIGN OF ADSORPTION COLUMN
This chapter provides the design of an adsorption column unit for the production pure glucose
syrup to be hydrogenated into sorbitol at atmospheric pressure (1atm). The adsorption column
unit is supposed to remove โadsorbโ colorants, impurities by the action of interpenetration of the
dextrose syrup and an activated carbon. The carbon can be activated by either steam or acid but
for the purpose of this project, steam activation of the carbon will be employed.
The unit will welcome only one incoming stream of dextrose syrup which has been produced by
heat liquefaction and saccharification process and the dextrose syrup will flow downstream to
the fixed bed in the column. The unit is required to operate at a temperature of 69-77oC to
enhance effective penetration and adsorption of impurities in the glucose syrup as it flows throw
the column.
5.2.1 ADSORPTION MECHANISM
The adsorption process is a mass transfer mechanism in which the mass in one medium adheres
to the surface of another medium. The substance being adsorbed is the adsorbate and the
adsorbing material is the adsorbent. The zone within which the adsorption takes place in the
adsorbent bed is known as the mass transfer zone (MTZ). The MTZ is also known as the
adsorption zone. According to (Eckhard 2012), the progress of the adsorption process can be
characterized by four consecutive steps;
First, transport of the adsorbate from the bulk liquid phase to the layer localized around the
adsorbent particle. As the liquid flows through the adsorbent bed, the adsorbates are transferred
to the localised area around the adsorbent particles. There is molecular diffusion of the adsorbate
in the fluid surrounding the adsorption granules.
62
After this, there is transport of the adsorbates through the boundary layer to the external surface
of the adsorbent.
Third, adsorbates move into the interior of the adsorbent particle (termed intraparticle diffusion
or internal diffusion) by diffusion in the pore liquid (pore diffusion) and/or by diffusion in the
adsorbed state along the internal surface (surface diffusion)
The energetic interaction between the adsorbate molecules is the last process and this will lead to
the final adsorption.
The adsorption process and its effectiveness is characterised by several factors such as the
surface area of the adsorbent, temperature, pressure effects, pH of the solution and isotherm
models. At the base of the column is a sieve of diameter less than the size of the adsorbent
particles. The sieve helps to filter the glucose as it exits the column. This is necessary to prevent
โsuspendedโ adsorbent particles from contaminating the glucose.
Regeneration of the bed becomes necessary when the adsorbent bed reaches its breakthrough
point. That is the bed has become exhausted. The regeneration of the bed can be done thermally
or by the use of acid. During the regeneration, the adsorbates are removed and the pores of the
bed reopen.
5.2.1.1 ADSORPTION EQUILIBRIUM
ISOTHERM MODELS CHARACTERIZING THE ADSORPTION PROCESS
To establish the correlation and to understand how the adsorbate is adsorbed by the adsorbent
(GAC) and to calculate the maximum adsorption capacity of the adsorbent bed, adsorption
isotherm must be established. Adsorption isotherm represents a relationship between the amount
of contaminant adsorbed per unit weight of carbon and its equilibrium concentration. The
63
amount of a substance that can be adsorbed on activated carbon depends on the nature of the
substance and its concentration, the surface structure of the activated carbon and the temperature
and pH of the solution.
There are many adsorption isotherm models but the commonly used models are the Freundlich
isotherm model, Langmuir isotherm and the BET (Brunauer, Emmett and Teller) isotherm.
Liquid-solid equilibrium between the concentrations of colorant, proteins (contaminants)
adsorbed on the carbon surface and the concentration of contaminants in the crude glucose is
described by the Langmuir isotherm. The Langmuir isotherm is the simplest and widely used
sorption isotherm(Wahyuningtyas et al. 2015), and occurs on the active site of monolayer
adsorbent surface, that is, adsorption involves the attachment of only one layer of molecules to
the surface.
The Langmuir isotherm is described by the equation below
๐ถ๐
๐๐=
1
๐๐๐พ1+
๐ถ๐
๐๐
๐๐ = ๐๐ ๐กโ๐ ๐๐๐ ๐ ๐๐ ๐๐๐ ๐๐๐๐๐ก๐ ๐๐๐ ๐๐๐๐๐ ๐๐๐ ๐ข๐๐๐ก ๐๐๐ ๐ ๐๐ ๐๐๐ ๐๐๐๐๐๐ก
๐ถ๐ = concentration of adsorbate remaining in the syrup after adsorption
๐พ1, ๐๐ are Langmuir constant representing the maximum adsorption capacity and energy of
adsorption respectively.
5.2.2 BREAKTHROUGH CURVE
When the fluid (glucose) is passed through the fixed adsorption bed, only the first part of the bed
work at adsorbing the solute until this part is saturated. Then the following section adsorb the
solute until it is, also, saturated. This trend continues until the whole bed is, saturated (usually
about 90% of saturation). Plotting the concentration in the effluent (incoming glucose syrup)
64
versus the volume of effluent (purified glucose syrup) gives the breakthrough curve. The Mass
Transfer Zone (MTZ) moves through the bed until it reached the lower part of the bed. At this
stage, the bead is said to be saturated and no adsorption can take place. There is the need to
regenerate the adsorbent bed either by steam or the use of acid.
5.2.3 BACKWASHING
Since the pressure loss increases with time due to the accumulation of particles in the adsorbent
bed, backwashing is necessary in certain time intervals. The backwashing is done by introducing
water at a moderate speed to flush out the adsorbed particles. This has to be done carefully to
prevent the mixing of the GAC particles.
5.2.4 COMPONENTS OF ADSORPTION COLUMN
All components of the equipment will be made of stainless steel except the carbon bed.
โข Column (Adsorber)
โข Adsorbent Bed (GAC)
โข Effluent Pipe
โข Influent Pipe
โข Metallic sieve
5.2.4.1 COLUMN (ADSORBER)
The column will be designed as closed pressure filters with circular cross section. The column
will be constructed with a food grade stainless steel. It houses the adsorbent bed and the metallic
sieve. The adsorbent in a fixed bed adsorber is located on a perforated metal (sieve) located at the
bottom of the adsorber, and the glucose streams downward through the adsorbent bed.
65
5.2.4.2 ADSORBENT BED DESIGN
FIXED-BEDS PROCESS PARAMETERS
The Column will consist of a single adsorbent fixed bed made from thermally activated carbon.
The adsorbent, which is the activated carbon granules in a fixed bed adsorber will be placed on a
perforated stainless steel sieve of the same diameter. The glucose syrup will flow downwards
through the adsorbent bed. The syrup will flow under gravity through the column. The volume of
the adsorber should account for adsorbed particles (adsorbates), bed expansion and maintenance,
thus, the activated carbon bed occupies approximately more than 66% of the bed.
Since the pressure loss increases with time due to the accumulation of particles in the adsorbent,
backwashing is necessary in certain time intervals to flush them out of the system.
5.2.4.2.1 ADSORBENT BED
Adsorbents are usually porous solids, and adsorption occurs mainly on the pore walls "inside"
particles. Adsorbents used in adsorption must be efficient to remove many and different
contaminants, have high adsorption capacity and rate of adsorption, and have high selectivity
for different concentration samples of adsorbents(Grassi et al. 2012). Examples of adsorbent
used in adsorption are listed below;
activated carbon (adsorbs organics)
silica gel (adsorbs moisture)
activated alumina (adsorbs moisture)
zeolites and molecular sieves
synthetic resins
66
Activated carbon is the adsorbent material selected for the purpose of this project. It is the vast
porous structure made from carbon that catches unwanted contaminants from the glucose syrup.
One gram of activated carbon will have a surface area equivalent to 1,000 square meters. This
permits the accumulation of a large number of contaminant molecules. Activated carbons come
in two forms, Powdered activated Carbon (PAC) and Granulated Activated Carbon (GAC). But
GAC is known to be used for adsorbent fixed beds. The GAC exhibits a crucial advantage over
the PAC due to its ability to be regenerated after it has been saturated with adsorbate. GAC
adsorbent of effective size ranges of 0.55-0.75mm will be used since it has been established that
smaller size of adsorbents gives better adsorption.
5.2.4.3 METALLIC SIEVE
The sieve will be made of a stainless steel and located at the bottom of the column. This
primarily filters the exiting glucose syrup. The adsorbent particles sometimes find their way into
the exiting glucose syrup and to the sieve help filter the syrup even as they exit the column.
5.2.4.4 INFLUENT AND EFFLUENT PIPES
The influent pipe introduces the glucose syrup into the adsorber from the top of the column and
the effluent pipe carries the refined syrup out of the column. The effluent pipe also serves as the
channel during backwashing.
67
5.2.3 ADSORPTION COLUMN DESIGN
ASSUMPTIONS
1. The adsorbent bed and column are cylindrical in shape
2. The diameter of the adsorbent bed is the same as the inside diameter of the adsorption
column.
3. The temperature of the column is maintained at 77oC.
Detail calculations are located at Appendix C
68
ADSORPTION COLUMN
69
GRANULATED ADSORBENT CARBON BED (GAC)
COLUMN LID
70
COLUMN BOTTOM
COLUMN SIEVE
71
Table 5.2 SUMMARIZED COLUMN PARAMETERS
Parameter Symbol Unit Calculated value
Volumetric flow rate ๐๐๐ ๐3/๐๐๐ฆ 3.3
Fixed-bed height ๐ป๐ต๐๐ ๐ 2
Fixed-bed diameter ๐ท๐ต๐๐ ๐ 1.5
Fixed-bed volume
๐๐ต๐๐ ๐ 3 3.6
Cross-sectional area
๐ด๐ต๐๐ ๐ 2 1.8
Empty Bed Contact
Time ๐๐๐ 42
๐ธ๐ต๐ถ๐ min 42
Effective contact
time
ฯ ๐๐๐ 12.5
Bed porosity ฮต - 0.3
Filtration Rate ๐น๐ m3
m2s
0.001
Adsorption column
volume
VAds ๐3 5.5
Adsorption column
Height
๐ป๐ด๐๐ ๐ 3
Void-filled volume VL ๐3 1.9
72
5.3 DESIGN OF A HYDROGENATION REACTOR
Hydrogenation is the most important unit operation in sorbitol production. Dextrose syrup is
converted to sorbitol by this unit operation. It is therefore necessary to ensure that optimal
conditions for temperature, pressure, feed to hydrogen gas flow and other factors are specified
accurately to obtain maximum sorbitol yield.
The objective of this chapter is to design a reactor for the efficient hydrogenation to obtain
sorbitol using the lowest possible hydrogen requirement, energy consumption and catalyst
consumption.
In designing the reactor, the choice of shape influences the holding capacity, the amount of
insulating materials required and the floor space requirement. Therefore, a vertically cylindrical
shape has been selected. The choice for the selection of the shape is;
Provides a relatively uniform area for insulation.
Provides a large holding capacity.
Less floor space requirement.
Easy to handle and automate
Economical in the use of energy.
5.3.1 OPERATING PRINCIPLES
Before operation, all air is extracted from the reactor. A 55%W of glucose solution continually
stirred to remove air bubbles and with a pH of about nine is mixed with hydrogen gas and
pumped through a heat exchanger to a temperature of about 100C.The liquid feed and hydrogen
gas are passed downward through the fixed bed of particulate high activity catalyst (Ru/Al). The
73
reactor is maintained at a temperature of 130โ, hydrogen pressure of 108.9 atm and a hydrogen
to feed ratio of 5:1. The product is collected above the reactor.
5.3.2 MAIN COMPONENTS OF THE REACTOR
Feed inlet: This section consists of an automated valve that opens when the feed is entering. It
has a diameter of 0.5 inches and is located at the bottom of the reactor to ensure that the feed
liquid droplets pass through the catalyst bed to achieve intimate contact with the catalyst.
Catalyst bed: The catalyst bed is made up of a layer of catalyst immersed on a carbon or silica.
It provides a medium for the contact of the reacting agent. The catalyst is coated with
monomolecular layer of carbon dioxide to prevent spontaneous oxidation of the highly active
catalyst when exposed to air during the loading into the reactor
Rotating stirrer: The rotating stirrer is a cylindrical rod with three impellers located on the
stirrer to enable efficient mixing of the product formed. It is made of stainless steel
Thermocouples: These are used to measure the temperature of the system to ensure that the
temperatures do not exceed the set point. They are located at the top, middle and bottom of the
reacting chamber.
5.3.3 MATERIALS FOR CONSTRUCTION
Certain factors are to be considered in designing equipment used in food processing. When
selecting materials for equipment design, some factors need to be considered. Prominent among
these are;
The ability to resist corrosion
The ease of fabrication.
The mechanical properties of the material
74
The ability to prevent product contaminations.
The cost of the material.
The most economical material that satisfies both process and mechanical requirements should be
selected for the design. Materials that give the lowest cost over the working life of the plant
allowing for maintenance and replacement.
The constructing materials for the reactor will be austenitic stainless steel (Type 304L), resin
bonded long fibre mineral wool, aluminum jacketing, cast iron and Rhethenium metal catalyst.
The reasons for selecting these materials are below.
AUSTENTIC STAINLESS STEEL (TYPE 304L)
1. Exhibit excellent corrosion resistance
2. High ease of fabrication
3. Easy to clean and weld.
4. Good mechanical properties
RHETHENIUM METAL CATALYST
1. Has high activity
2. High selectivity
3. Fast filtration rate
4. Can be recycled
75
CAST IRON
1. Relatively low melting point.
2. Good fluidity and castability.
3. Resistance to deformation and wear.
4. Excellent machinability.
5.3.4 SUMMARISED TABLE OF PARAMETERS OF THE HYDROGENATION REACTOR
PARAMETER SYMBOL UNIT VALUE
Volume of the hydrogenation
tank
m2 10.2
Volume of the reacting
chamber
m2 7.64
Number of impellers - 3
Impeller speed m/s 13.2
Synchronous speed Rpm 240
Volume of catalyst bed m3 6.74
Area of catalyst on bed m2 33.7
76
CHAPTER SIX
6 PROCESS CONTROL
Process control in the sorbitol processing plant
For the sorbitol production plant, process control is a major factor. All equipment in the section
of this plant need a control process to ensure that the same quality of product is obtained at each
production. The table below talks of the equipment that need controls and the type of controls
that are needed
Equipment Control
Starch plant 1. Pressure
2. Temperature
3. Fluid flow
4. Level control
Adsorption column 1. Pressure
2. Temperature
3. Fluid flow
Liquefaction vessel 4. Pressure
5. Temperature
6. Fluid flow
7. Level control
Hydrogenation reactor 8. Pressure
9. Temperature
77
10. Fluid flow
Vacuum evaporator 1. Pressure
2. Temperature
Glucose storage tank 1. Fluid flow
Cabinet Dryer 1. Temperature
Packaging machine 1. Weight measure
78
CHAPTER SEVEN
7.0 GENERAL CONSIDERATION
Cassava (Manihot esculenta), is perennial crop which is predominantly grown in the tropics
primarily for its starchy roots, though its high protein leaves are consumed by some communities
as vegetables. Cassava constitutes 22 percent of Ghanaโs agricultural GDP and one of Ghanaโs
main staple crops with an annual production above 10 million metric tonnes in the last decade.
Also, in terms of area harvested, cassava is now the second largest crop as it has been recently
superseded by maize. The crop ranks first among its root and tuber contemporaries inโterms of
production (13,504 MT in 2010) and has an estimated per capita consumption of 152.9 kg/year
(MOFA, 2010 and FAOSTAT, 2013). Ghana is the 6th world producer of cassava in terms of
value. Ghanaโs ranking remained unchanged over the period of analysis 2005-2010 (FAOSTAT,
2013). The crop is widely valued as a low cost carbohydrate source for urban consumers
(Hillocks, 2002) and the food security that it provides (Siritunga and Sayre, 2003).
The carbohydrate values present in cassava are consistent which range from 32% to 35% on
fresh basis and 80% to 90% on dry matter basis (Montagnac et al., 2009; Zvinavashe et al.,
2011). Cassava has a high moisture content that ranges from 33.14% to 45.86% for the local
cultivars. (Afoakwa et al.,2012). Cassava contains about 1-2% protein as compared to maize and
sorghum which have about 10g protein/100 g fresh weights (Charles et al., 2005; Montagnac,
2009), which makes it suitable for sorbitol production. Cyanide is the most toxic factor
restricting the consumption and utilization of cassava roots and leaves, but the level of
cyanogenic glucosides can be controlled through specific water treatments (Meridian Institute,
2009).
79
The government of Ghana in its efforts to develop starch production from cassava, implemented
a specific cassava policy in the Presidential Special Initiative (PSI). The PSI in Ghana is one of
similar initiatives undertaken in other cassava producing countries, and supported by the
NEPADโs, Pan-African Cassava Initiative (NPACI) launched in January 2004. The policy has
given birth to the Ayensu cassava starch processing plant.
7.1. SITE SELECTION
Our plant will be located at Asamankese in the Eastern Region of Ghana where about 27% of the
total national cassava production is obtained. The town also provides all the other requirements
mentioned below.The location of a plant plays an important role in the optimization of various
process factors. The following factors were considered in selecting the site.
โข Availability of raw material
โข Cheap labour.
โข Nearness of plant to target market
โข Availability of electricity and water
โข Availability of land size needed for plant construction
The total land area of 54,000ft2 is required for the plant and the number of plots equivalent is
eight (8).
80
Pie chart displaying the distribution of cassava production in Ghana
81
7.2 SANITATION DESIGN CONSIDERATIONS
Some sanitary design features include the following.
1. All parts of the equipment are made of impervious material which prevents retention of
fluids that can create a good atmosphere for microbial growth.
2. All surfaces that come into contact with the feed are made of food grade stainless steel.
3. All surfaces are smooth, continuous and without crevices or cracks, thus making cleaning
and maintenance less time consuming.
7.3 SAFETY CONSIDERATIONS IN THE PLANT
There are hydrogen detectors at all parts of the plant to ensure that any hydrogen gas leakage is
identified and controlled. This is very critical for the safety of the workers and the production in
general. Signs will be placed at all points to inform workers on where they must pass and avoid
in the plant to minimize accidents.
The product at any stage of the plant is passed through a metal detector to identify any metal that
might have been introduced during the processing period.
7.4 GIPC LAWS IN RELATION TO INVESTMENT, CAPITAL AND LABOR
EMPLOYMENT
According to the Ghana Investment Promotion Centre Act, 2013, (Act 865), enterprises which
are wholly owned by a Ghanaian after being registered or incorporated to the Centre will have
some benefits and incentives that can enable them to trade effectively in the country.
An enterprise registered by the centre is entitled to all benefits under Internal Revenue Act,2000
(Act 529), Value added Tax Act,1998 (Act 564) and chapters 82,84,85 and 98 of the Customs
82
Harmonized Commodity and Tarrif code schedule to the Customs, Excise and Preventive
Services(management) Act,1993 (PNDCL 330).
In relation to investment guarantees, transfer of capital, profits, dividends and personal
remittances, the enterprise shall through an authorized bank be guaranteed unconditional
transferability in freely convertible currency of dividends or net profit, payment in respect to
technology transfer agreement, remittance of proceeds, net of all taxes and other obligation in the
event of the sale or liquidation of the enterprise or any attribute to the enterprise.
On labor and employment, an enterprise registered under this Centre, shall abide by the
application of labor legislature. The labor legislations may be agreements made between
employees but the agreement shall not establish standards lower than the mandatory
requirements under the laws of Ghana.
7.5 UTILITIES
The main utilities used in the plant are water and electricity. The Municipal water will be the
main source of water supply. All water to be used in the plant will be treated before usage. There
will be a recycling of waste water before discharging into major drains.
The Electricity Company will also be the main source of electricity supply. There will be a
stand-by generator to power the plant in case of power outages.
83
CHAPTER EIGHT
8.0 ECONOMIC ANALYSIS
Economic analysis is an essential part of plant design and development. Economic analysis
provides a tool for determining the viability of a project. For a project to be acceptable, the plant
design must present a process that is capable of operating under conditions which will yield a
profit. The economic analysis therefore enables the food process engineer to come out with the
preliminary information of the money flow in the project operations before proceeding to erect
the plant. Evaluation of costs in the preliminary design phases of a plant is sometimes called
โguesstimationโ but the appropriate designation is predesign cost estimation. The predesign cost
estimation include the capital investment, which is a sum of the fixed-capital investment for
physical equipment and facilities in the plant plus working capital which must be available to
pay salaries, keep raw materials and products on hand, and handle other special items requiring a
direct cash outlay (Peters and Timmerhaus, 1991). The predesign cost is needed to help top
management of a company or the funders of a project make an informed decision as to whether
to proceed to invest in a project or not.
Economic analysis shows the profitability and viability of a project. To do this effectively and
accurately, the analysis has been divided into three major categories to enable the food process
engineer make a sound guess of the viability of the project;
A. Total Capital Investment (TCI)
B. Total Production Cost (TPC)
C. Profitability Analysis (PA)
84
BASIS FOR ECONOMIC ANALYSIS
Start construction: June 2016
Complete of construction: November 2017
Commencement of production: December 2017
Working periods: 300 days/year
Production rate of sorbitol 3,000 kg/day
Plant life: 20 years
Exchange rate: ๐๐๐ท 1 = 3.84 ๐บ๐ป๐ถ (BOG rates, retrieved on 2/4/2016)
Interest rate: 35 %
Tax rate: 12.5%
8.1 TOTAL CAPITAL INVESTMENT
Total capital investment is the money needed to purchase and install the necessary machinery
and equipment, and for the operations of the plant. The total capital investment is subdivided into
fixed capital investment and working capital.
๐ถ๐ = ๐ถ๐น + ๐ถ๐
where
๐ถ๐ = total capital investment
๐ถ๐น = fixed capital investment
๐ถ๐ =working capital
85
Fixed capital investment is the money needed to purchase and install the necessary machinery
and equipment, and plant facilities.
Working capital is necessary for the operation of the plant. It is usually estimated as a fraction of
the fixed capital: ๐ถ๐ = ๐ถ๐น ร ๐๐
Where ๐๐ is the ratio of working to fixed capital. The ratio fw, varies from 10-20% in most
process industries, but it may be as high as 50% in plants processing products of seasonal
demand, or seasonal production of raw materials, e.g., fruits and vegetables. The following value
is suggested for food industries: ๐๐ =0.25 (Maroulis and Saravacos, 2003).
8.2 ESTIMATION OF FIXED CAPITAL INVESTMENT
8.2.1 EQUIPMENT COST
When costing equipment for preliminary design, a number empirical methods and rules are used
to make quick and accurate approximation of the cost. For instance, when the current cost of
equipment is not known because it has a different capacity from the off-the-shelf equipment, the
cost of that equipment can be approximated using the six-tenths-factor rule by knowing the old
cost and capacity of the off-the-shelf equipment as well as the desired capacity of the equipment
(Chilton, 1960). The formula is stated below;
๐ถ = ๐ถ๐ (๐
๐๐)๐
โฆโฆโฆโฆโฆโฆโฆโฆ . .8.1
Where;
C = cost of equipment
86
๐ถ๐=cost of off-the-shelf equipment
M = capacity of equipment
๐๐= capacity of off-the-shelf equipment
๐= scale index. It is different for the different equipment in the starch factory.
Considering the time value of money, inflation and rise in prices of materials of production, the
cost of the equipment can be brought up-to-date and put on a common basis using the Marshall
and Smith index, which can be calculated as;
๐๐๐๐ โ๐๐๐ & ๐๐ค๐๐๐ก (๐&๐)๐๐๐๐๐ฅ = 1100 + 20(๐ฆ๐๐๐ โ 2000)โฆโฆโฆโฆโฆโฆโฆโฆ8.2
All the cost data of imported equipment for the starch factory is from Zhengzhou Bizoe Import &
Export Trading Co. Ltd., China. And they are all declared free on board (FOB). The cost of
imported equipment for the rest of the equipment for other sections of the plant were obtained
from Alibaba.
87
8.2.2 TOTAL IMPORTED EQUIPMENT COST
Table 8.1 Cost of imported equipment
EQUIPMENT MATERIAL OF
CONSTRUCTION
QUANTITY UNIT COST
(USD)
TOTAL COST
(USD)
STARCH PLANT
Scraper lifter CS 1 7,500 7,500
Washing peeler CS & SS 1 8,500 8,500
Paddle wash machine CS & SS 1 9,000 9,000
Inclined squirrel cage
washing machine
CS 1 9,500 9,500
Curved mesh crusher SS 1 12,500 12,500
Slurry pump SS 6 325 1,950
Hydro cyclone SS 12 sets 58,000 58,000
Grit catcher SS 1 1,000 1,000
Starch separator SS 1 10,500 10,500
Vacuum dehydrator SS 1 10,500 10,500
Inspection table Plastic belt 1 1,500 1,500
Electromagnetic flow
meter
- 2 2,500 5,000
Sub-total USD 135,450
88
GLUCOSE PREPARATION PLANT
Equipment MATERIAL OF
CONSTRUCTION
Quantity Unit price (USD) Total Price (USD)
CSRT (liquefaction) SS 1 6,500 6,500
CSRT(saccharification) SS 1 6,000 6,000
Acidification Tank CS 1 4,000 4,000
Plate Heat Exchanger CS 1 1,250 1,250
Adsorption Column SS 1 71,407 71,407
Ion Exchange Column
(Cation)
SS 1 10,000 10,000
Ion Exchange Column
(Anion)
SS 1 10,000 10,000
Dosing tank SS 2 1,000 2,000
Centrifugal pump CS 4 400 1,600
In-Line Pumps CS 4 300 1,200
Sub-total USD 113,957
SORBITOL PREPARATION PLANT
EQUIPMENT MATERIAL OF
CONSTRUCTION
QUANTITY UNIT COST
(USD)
TOTAL COST
(USD)
Hydrogenation reactor SS 1 28,000 28,000
Evaporator SS 1 10,000 10000
89
Crystalliser SS 1 11750 11750
Pressure filter SS 1 9,000 9000
Roller mill SS 1 500 500
Cabinet dryer CS 1 8,000 8000
Conveyor belts CS 3 500 1500
Heat exchanger SS 1 1,250 1250
Pumps SS 5 400 2000
Packaging equipment SS 1 960 960
Sub- total USD 72965
TOTAL USD ๐๐๐, ๐๐๐
Total cost of imported equipment
= starch production equipment cost + glucose preparation equipment cost
+ ๐ ๐๐๐๐๐ก๐๐ production plant equipment cost
๐ก๐๐ก๐๐ ๐๐๐๐๐๐ก๐๐ ๐๐๐ข๐๐๐๐๐๐ก ๐๐๐ ๐ก = 135,450 + 113,957 + 72,965
= ๐ผ๐บ๐ซ ๐๐๐, ๐๐๐
Free On Board (FOB) = freight charges + total cost of imported equipment
But freight charges are 2% of total cost of imported equipment
๐๐๐ก๐๐ ๐๐๐ ๐ก ๐๐ ๐๐๐๐๐๐ก๐๐ ๐๐๐ข๐๐๐๐๐๐ก = 1.02 ร 322,372
= ๐ผ๐บ๐ซ ๐๐๐, ๐๐๐
90
The total cost of imported processing equipment is USD 328,819 and considering the current
exchange rate of USD1= GHศป 3.84, the Ghanaian cedi equivalent is GHศป 1,262,667
8.2.3 COST OF LOCALLY ACQUIRED AND FABRICATED EQUIPMENT
Table 8.2. Cost of locally acquired and fabricated equipment and vehicles
ITEM MATERIAL OF
CONSTRUCTION
QUANTITY UNIT COST
(GHศป)
TOTAL
COST (GHศป)
Sand removing rotary
drum
CS 1 19,300 19,300
Mixing tank SS 2 2,400 4,800
Cake Collecting tray Plastics 10 10 100
Storage tank SS 2 4,000 8,000
pick-ups 1 45,000 45,000
Mercedes-Benz Ateco,
816
1 55,000 55,000
TOTAL GHยข 132,200.00
Total cost of equipment = cost of imported equipment + cost of locally acquired equipment
= 1,262,667 + 132,200.00
= ๐ฎ๐ฏยข๐, ๐๐๐, ๐๐๐
91
Estimates;
Spare parts = 10% of total cost of equipment
Equipment spare parts cost = 0.10 ร 1,394,867
= ๐ฎ๐ฏยข 139,487
Handling and transportation = 0.5% of total equipment cost
= 0.05 ร 1,394,867
= ๐ฎ๐ฏยข ๐๐, ๐๐๐
8.2.4 TOTAL PURCHASED EQUIPMENT COST (TPE)
Table 8.3 Total Purchase Equipment Cost
COMPONENT ESTIMATED COST (GHยข)
Equipment Cost 1,394,867
Spare parts 139487
Handling and transportation 69,743
Total TPE GHยข 1,604,097
92
Table 8.4 Estimation of Direct And Indirect Cost (Peter and Timmerhaus approach, 1991)
DIRECT COST
ITEM COST FACTOR OF TPE ESTIMATED COST (GHศป)
Total Purchased equipment (TPE) 1.0 1,604,097
Equipment installation 0.47 753,926
Piping 0.15 240,615
Electrical installation 0.10 160,410
Building/ auxiliary 0.20 320,819
Service and land improvement 0.20 320,819
Instrumentation control 0.20 320,819
Total direct cost (DC) = GHศป 3,721,506
INDIRECT COST COST FACTOR OF DC ESTIMATED COST
Engineering and supervision 0.10 372,151
Construction expense and
contractor
0.12 446,581
Contingency 0.05 186,075
Total indirect cost GHศป 1,004,807
93
Total Cost of Equipment = Direct Cost + Indirect Cost
Total Cost of Equipment = GHศป 4,726,314
8.2.3 COST OF LAND
Cost of Land (70ft by 100ft) at Asamankese in the Eastern Region is GHศป 3,500.
Considering a total land area of 54,000 ft2, the number of plots required is eight (8)
Therefore, the total cost of the five plots = 8 ร 3500
= ๐๐ศป ๐๐, ๐๐๐
8.3.3 ESTIMATION OF FIXED CAPITAL INVESTMENT
๐น๐๐ฅ๐๐ ๐๐๐๐๐ก๐๐ ๐๐๐ฃ๐๐ ๐ก๐๐๐๐ก(๐ถ๐น) = ๐ก๐๐ก๐๐ ๐๐๐ ๐ก ๐๐ ๐๐๐ข๐๐๐๐๐๐ก + ๐๐๐ ๐ก ๐๐ ๐๐๐๐
๐ถ๐น = 4,426,314 + 28,000
๐๐ศป ๐, ๐๐๐, ๐๐๐
8.3 ESTIMATION OF WORKING CAPITAL
The working capital ๐ถ๐ค consists of the total amount of money invested in raw materials and
supplies carried in stock, finished and semi-finished products, accounts receivable and payable,
and cash kept on hand. It is usually estimated as a fraction of the fixed capital:
๐ถ๐ = ๐๐ ร ๐ถ๐
94
According to Peter and Timmerhaus (1991) and Zacharia and George (2003), the ratio of
working to fixed capital fw varies from 10-20% in most process industries, but for food
industries, it is estimated to be 15% of the total capital investment (๐ถ๐);
๐ถ๐ = ๐ถ๐น + ๐ถ๐
๐ถ๐ = 0.15๐ถ๐
๐ถ๐ = ๐ถ๐น + 0.15๐ถ๐
๐ถ๐น = ๐ถ๐ โ 0.15๐ถ๐
= 0.85๐ถ๐
๐ถ๐ =๐ถ๐น
0.85=
4,754,314
0.85
= ๐ฎ๐ฏ๐ช ๐, ๐๐๐, ๐๐๐
Hence ๐ถ๐ = 0.15 ร 5,593,311
= ๐ฎ๐ฏ๐ช ๐๐๐, ๐๐๐
SUMMARY OF INVESTMENT COST
Total Capital Investment (๐๐) GHศป ๐, ๐๐๐, ๐๐๐
Working Capital (๐๐) GHศป 838,997
Fixed Capital Investment (๐๐ ) GHศป ๐, ๐๐๐, ๐๐๐
95
8.4 ESTIMATION OF TOTAL PRODUCTION COST (TPC)
Another equally important part of the economic analysis is the estimation of costs for operating
the plant and selling the products. According to Maroulis and Saravacos (2003), the total
production cost (TPC) is the sum of the manufacturing cost and the non-manufacturing cost or
general expenses. The equations below show the relationship between total production cost
(TPC), manufacturing cost and general expense, as presented by Peter and Timmerhaus (1991)
Total Production Cost(TPC) = Manufacturing Cost(๐๐) + General Expenses (๐๐)
Manufacturing Cost(๐๐) = Direct Production Cost + Fixed Charges + Plant Overhead Cost
General Expenses (๐๐) = Administration Expenses + Distribution and Marketing Expenses
8.4.1 ESTIMATION OF MANUFACTURING COST (๐๐)
Manufacturing costs are also known as operating or productions costs and are divided into three
major categories; direct production costs, fixed charges and plant overhead costs.
8.4.2 DIRECT PRODUCTION COST
According to Peter and Timmerhaus (1991) the direct production cost include expenses directly
associated with the manufacturing operation. This type of cost involves expenditures for raw
materials, direct operating labor; supervisory and clerical labor directly connected with the
manufacturing operation; plant maintenance and repairs; operating supplies, power, utilities and
catalysts.
96
Table 8.5. Estimation Of Direct Production Cost
ITEM ANNUAL QUANTITY
(Kg)
COST PER UNIT
(GHศป/Kg)
TOTAL (GHศป)
RAW MATERIAL
cassava 3,000,000 0.30 900,000
Sulphur Dioxide 4,800 0.39 1,872
Alpha Amylase 2,615 135 353,079
Gluco-amylase 2,610 135 352,350
Sulphuric Acid 65,400 5 327,000
Hydrogen 300,000 1.92 576,000
Catalyst 450 230 103,500
RAW MATERIAL TOTAL COST GHศป 2,613,801
OPERATION SUPPLIES AND UTILITIES
Starch
production
Dextrose
production
Sorbitol production
Power(kWh/year) 144,240 41,400 43,546
Cost of power
(GHศป)
139611.19 40,485.00 42,148.00
Water (Litres/year) 24,000,000 1,380,000 514,296
Total power consumption per year: 245,599 kWh
97
According to Electricity Company of Ghana tariffs for industries, 96.79 GHp/KWh
Hence the total cost of power consumption
= GHศป 222,244.00
Total water used per year; 25,894,296 liters
According to Ghana water company tariffs for industries, 380.0075GHp/1000litres
So total cost of water consumption is GHศป 98,400.00
Total utilities cost = cost of power consumed + cost of water used
= 222,244 + 98,400
= ๐ฎ๐ฏ๐ช ๐๐๐, ๐๐๐
COST OF PACKAGING MATERIALS
Number of card boxes needed daily = 60
Number of required boxes annually = 18,000
Unit cost of corrugated board =GHศป 3.00
Total cost of board annually required = GHศป 54,000
Unit cost of low density polyethylene = GHศป 0.80
Annual cost of low density polyethylene =GHศป 14,400
Hence total cost of packaging material annually =GHศป 68,400
98
Table 8.6. Estimation of Operation Labour
ITEM QUANTITY Monthly salary
(GHศป)
TOTAL annual
salary (GHศป)
(A) MANAGEMENT
General Manager 1 2,500 30,000
Starch plant manager 1 1,500 18,000
Glucose production plant manager 1 1,500 18,000
Sorbitol preparation plant manger 1 1,500 18,000
Accountant 1 1,500 18,000
Human resource manager 1 1,800 21,600
Mechanical superintendent 1 1,300 15,600
Technical supervisor 1 1,300 15,600
Internal auditor 1 1,000 12,000
(B) Indirect labour
Agent for root supply 1 500 6,000
Office of clerk 1 825 9,900
Guards 2 500 12,000
Cleaners 3 300 10,800
99
(C) Direct labour
Processing technician 3 3,000 36,000
Quality-control technician 1 1000 12,000
Skilled labour for processing
operations
7 600 42,000
Skilled workers (drivers, electrical
maintenance)
5 600 30,000
Unskilled workers for processing 4 300 14,400
Total labour cost GHศป 329,500
Social security =12.5% of total labour cost
= 0.125 ร 329,500
= ๐บ๐ป๐ถ 41,188
Total operating labour cost = 329,500 + 41,188
= GHศป 370,688.00
100
FIXED CHARGES
These are expenses which remain somewhat constant from year to year and do not vary
substantially with changes in production rate (Peter and Timmerhaus, 1991). Some of these fixed
charges include depreciation, taxes, insurance and rent.
DEPRECIATION COST
Depreciation is defined by Perry and Green (2008) as the lose value of an asset due to physical
deterioration, technological advances, economic changes among similar factors. Several methods
are employed in the determination of the depreciation cost; however, in this project the straight
line method is applied.
๐ =๐ถ๐น โ ๐๐
๐
d = annual depreciation (GHศป/yr) =
๐ถ๐น = the Initial Fixed Capital Investment = GHศป 4,754,314
๐๐ = Salvage value = 10% ๐ถ๐น = 0.10(4,754,314) = GHศป 475,431
๐ = Service life or useful life of plant (yrs) = 20yrs
๐ =4,754,314 โ 475,431
20
= ๐ฎ๐ฏ๐ช ๐๐๐, ๐๐๐
101
INSURANCE
Insurance amounts to 1% of the fixed capital investment.
๐ผ๐๐ ๐ข๐๐๐๐๐ = 0.01 ร ๐ถ๐น
= 0.01 ร4,754,314
= ๐๐๐ ๐๐, ๐๐๐
๐ ๐ข๐ฑ๐๐ ๐๐ก๐๐ซ๐ ๐๐ฌ = ๐๐๐ฉ๐ซ๐๐๐ข๐๐ญ๐ข๐จ๐ง + ๐๐ง๐ฌ๐ฎ๐ซ๐๐ง๐๐
๐ญ๐๐๐๐ ๐ช๐๐๐๐๐๐ = 213,944 + 47,543
= ๐ฎ๐ฏ๐ช 261,487
PLANT OVERHEAD COST
Plant-overhead costs are for hospital and medical services; general plant maintenance and
overhead; safety services; payroll overhead including pensions, vacation allowances, social
security, and life insurance; packaging, restaurant and recreation facilities, salvage services,
control laboratories, property protection, plant superintendence, warehouse and storage facilities,
and special employee benefits. It is estimated to be 15% of total production cost (Peter and
Timmerhaus, 1991).
102
Table 8.7. SUMMARY OF MANUFACTURING COST (๐๐)
COMPONENT COST FACTOR TOTAL COST(GHศป)
DIRECT PRODUCTION COST
Raw Material - 2,613,801
Utilities (Electricity and Water) - 320,644
Operation Labour Cost - 370,688
Operating supervision 15% of operating labour 55,603
Maintenance and repairs 6% of ๐ถ๐น 285,258
Laboratory charges 10% of operating labour 37,069
Total Direct Production Cost GHศป 3,313,433
FIXED CHARGES
Depreciation 213,944
Insurance 47,543
Local Taxes 3% of ๐ถ๐น 142,629
Total Fixed Charges GHศป 404,116
PLANT OVERHEAD COST
Plant Overhead Cost 15% TPC 0.15TPC
Total Manufacturing Cost (Cm )= 3,717,549 + 0.15TPC
103
Table 8.8. SUMMARY OF GENERAL EXPENSES (๐๐)
COMPONENT COST FACTOR TOTAL COST
Sales and distribution expenses 2% of TPC 0.02TPC
Research and development 2% of TPC 0.02TPC
Administrative costs 2% of TPC 0.02TPC
TOTAL (GE) 0.06TPC
Total Production Cost(TPC) = Manufacturing Cost(๐๐) + General Expenses (๐๐)
TPC = ๐๐ + ๐๐
TPC = 3,717,549 + 0.15TPC + 0.06TPC
TPC = 3,717,549 + 0.21TPC
TPC โ 0.21TPC = 3,717,549
TPC(1 โ 0.21) = 3,717,549
TPC =3,717,549
0.79
๐๐๐ = ๐๐ศป ๐, ๐๐๐, ๐๐๐
104
Plant overhead = 0.15 ร TPC
= GHศป 705,864
General Expenses = 0.06TPC
= 0.06 ร 4,705,758
= GHศป 282,345
Total Manufacturing Cost = 3,717,549 + 0.15TPC
= 3,717,549 + 0.15 x 4,705,758
The Total Manufacturing Cost = GHศป 4,423,413
SUMMARY OF TOTAL PRODUCTION COST
DIRECT PRODUCTION COST GHศป 3,313,433
FIXED CHARGES GHศป 404,116
PLANT OVERHEAD COST GHศป 705,864
GENERAL EXPENSES GHศป 282,345
MANUFACTURING COST (๐๐) GHศป 4,423,413
105
8.5 PROFITABILITY ANALYSIS
The word profitability is used as the general term for the measure of the amount of profit that can
be obtained from a given situation. Profitability analysis therefore attempts to measure the
attractiveness of the project in comparison to other competing investments. In other words, it
enables an investor to know the risk involved in investing a project, and also know what to
expect during the operation of the project.
Profitability analysis is done to measure the amount of profit that can be obtained from a given
situation.
The profitability of this project will be evaluated using the following common methods
Rate of Return on Investment (ROI)
Internal Rate of Return (IRR) or the discounted cash flow based on full life performance
90
Net present value (NPV)
Capitalized cost
Payback period or pay out period (PBP)
8.5.1 ESTIMATION OF ANNUAL REVENUE
Calculating the selling price
Using the mark-up method to estimate the selling price of the sorbitol
๐๐๐๐๐๐๐ ๐๐๐๐๐ = ๐๐๐ก๐๐ ๐ถ๐๐ ๐ก ร (1 + ๐๐๐๐ โ ๐ข๐ ๐๐๐๐๐๐๐ก)
Production rate of sorbitol = 3,000 kg per day
106
Total annual production = 3,000๐๐
๐๐๐ฆ ร 300
๐๐๐ฆ
๐ฆ๐๐๐
= 900,000 ๐๐
Therefore, the total annual production is 900,000 ๐๐
Number of 50 kg bags per year = 900,000
50
= 18,000 ๐๐๐ฅ๐๐
Total Production Cost (TPC) = GHศป 4,705,758
The cost of producing a 50 kg box =๐๐๐ถ
Number of 50 kg bags per year
= 4,705,758
18,000
= ๐บ๐ป๐ถ 261.
Using a total mark-up percent of 65 %
๐๐๐๐๐๐๐ ๐๐๐๐๐ = ๐๐๐ก๐๐ ๐ถ๐๐ ๐ก ร (1 + ๐๐๐๐ โ ๐ข๐ ๐๐๐๐๐๐๐ก)
= 261 ร (1 + 0.65)
= GHศป 431
Therefore, a 50 Kg box of sorbitol will be sold for GHศป430
Revenue at the end of production year = ๐๐๐๐๐๐๐ ๐๐๐๐๐ ร No. of 50 kg bags produced annually
= 430 ร 18,000
= ๐บ๐ป๐ถ 7,740,000
Hence, the Revenue at the end of production year is ๐บ๐ป๐ถ 7,740,000
107
DATA
Fixed Capital Investment (CF) = GHศป 4,754,314
Working Capital (Cw) =GHศป 838,997
Total Capital Investment (CT) = GHศป 5,593,311
Total Production Cost (TPC) = GHศป 4,705,758
Manufacturing Cost (CM) = GHศป 4,423,413
Salvage Value (VS) = GHศป 475,431
Depreciation (D) =GHศป 213,944
ECONOMIC ENVIRONMENT
Depreciation coefficient ๐ ๐ =๐
๐=
1
20= 0.05
Depreciation d = GH ศป 213,944 per year
Interest rate, i (discount rate) = 35 %
Tax rate, t =12.50%
Plant life, n= 20 years
Capital recovery, ๐ =๐(1+๐)๐
(1+๐)๐โ1
= 0.35(1+0.35)20
(1+0.35)20โ1 =0.35
108
8.5.2 GROSS PROFIT (PG)
๐บ๐๐๐ ๐ ๐๐๐๐๐๐ก = ๐ ๐๐ฃ๐๐๐ข๐ โ ๐ถ๐๐ ๐ก ๐๐ ๐๐๐๐ข๐๐๐๐ก๐ข๐๐๐๐
PG = 7,740,000 โ4,423,413
PG = GH ศป 3,316,587
8.5.3 TAXABLE INCOME (R)
๐๐๐ฅ๐๐๐๐ ๐ผ๐๐๐๐๐ = Gross Profit โ ๐๐๐๐๐๐๐๐๐ก๐๐๐ ๐๐๐๐๐๐๐๐๐๐๐ก ร Fixed Capital Investment
๐ = ๐๐บ โ ๐ ๐ ร ๐ถ๐น
= 3,316,587 โ 0.05 ร 4,754,314
= ๐บ๐ปศป 3,078,871
8.5.4 INCOME TAX (T)
๐ผ๐๐๐๐๐ ๐๐๐ฅ = Taxable Income ร Tax rate
๐ = ๐ ร ๐ก
= 3,078,871 ร 0.125
= ๐บ๐ปศป 348,859
8.5.5 ANNUAL PROFIT AFTER TAX (P)
๐ด๐๐๐ข๐๐ ๐๐๐๐๐๐ก ๐๐๐ก๐๐ ๐๐๐ฅ = ๐บ๐๐๐ ๐ ๐๐๐๐๐๐ก โ ๐ผ๐๐๐๐๐ ๐๐๐ฅ
P = PG โ T
= 3,316,587 โ 348,859
= GH ศป 2,931,728
109
8.5.6 ANNUAL CASH FLOW (CF)
Annual Cash Flow = Annual Profit after Tax + Depreciation
๐ถ๐น = ๐ + ๐
= 2,931,728+ 213,944
= GH ศป 3,145,672
8.5.7 NET PROFIT (PN)
Net Profit = Annual Cash Flow + Recovery factor x Fixed Capital Investment
PN= CF - e๐ถ๐น
= 3,145,672โ 0.35 ร 4,754,314
= GH ศป 2,907,956
110
8.6 FINANCIAL APPRAISAL
According to Peter and Timmerhaus, the following are used to determine the profitability of a
project
1. Payback Period (PBP)
2. Discounted payback period (DPB)
3. Return on investment (ROI)
4. Net Present Value (NPV)
5. Cumulative Cash flow (CCF)
8.6.1 PAYBACK PERIOD (PBP)
Payback Period=๐๐๐ก๐๐ ๐ถ๐๐๐๐ก๐๐ ๐ผ๐๐ฃ๐๐ ๐ก๐๐๐๐ก
๐ด๐๐๐ข๐๐ ๐๐๐๐๐๐ก ๐๐๐ก๐๐ ๐๐๐ฅ
๐๐ต๐ =๐ถ๐
๐
=5,593,311
2,931,728 = 1.9 ๐ฆ๐๐๐๐
8.6.2 DISCOUNTED PAYBACK PERIOD
DPB =
[ ln (
11 โ i(PBP)
)
ln (1 + i)
]
=
[ ln (
11 โ 0.35(1.9)
)
ln (1 + 0.35)
]
๐๐๐ = ๐. ๐ ๐ฒ๐๐๐ซ๐ฌ
111
8.6.3 CUMULATIVE CASH FLOW
CCF = โCT + (n x CF)
CCF = โ5,593,311 + (20 x 3,145,672)
CCF = GHศป 57,320,129
8.7 SENSITIVITY ANALYSIS
Return on Investment (ROI):
It is expressed on annual percentage basis:
๐ ๐๐ก๐ข๐๐ ๐๐ ๐๐๐ฃ๐๐ ๐ก๐๐๐๐ก = ๐
๐ถ๐ร 100%
ROI = 2,931,728
5,593,311 ร 100%
= 52.4%
Therefore, the Return on Investment (ROI) is 52.4%
Net Present Value (NPV)
Capital Recovery Factor, e = 0.35
NPV = [โCT +P
e]
NPV = [โ5,593,311 +2,931,728
0.35]
NPV = GHศป 2,783,058
Profitability Index (PI)
PI = NPV
CT=
2,783,058
5,593,311
PI = 0.5
Since the Profitability Index is 0.5 which is greater than zero, the project is viable.
112
8.7.1 INTERNAL RATE OF RETURNS (IRR)
Internal rate of return is a discount rate that makes the net present value (NPV) of all cash flows
from a particular project equal to zero
Table 8.9 A table of interest rate against Net Present Value
Interest Rate, i Net Present Value (NPV)
0.1 7903128.5
0.2 7843035.81
0.3 3555883
0.4 1296004.68
0.5 -77450.17
0.6 -995841
0.7 -1652622.4
0.8 -2145208.5
0.9 -2528331
113
114
CHAPTER NINE
9.0 CONCLUSION AND RECOMMENDATION
The plant design for the production of sorbitol form cassava is viable and lucrative, if completed
successfully since there is increasing demand for sorbitol and also cassava the raw material is
readily available. The project will impact the lives of Ghanaians by providing market for cassava
farmers and also creates jobs for Ghanaians.
We recommend that an alcohol distillation plant be attached to the sorbitol plant to process the
output liquor from the sorbitol filtration unit into alcohol.
115
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119
APPENDIX A
MATERIAL BALANCE MASS BALANCE FOR THE STARCH EXTRACTION PLANT
SORTING AND SAND REMOVING
ASUMPTIONS
The plant capacity is 15 tonnes of cassava per day
The amount of cassava sorted as waste is 1% of the total cassava received
The mass of water used for the washing the roots is 141.4% of the mass of cassava that
passed the sorting process
The peel constitutes 1.5% of the total root mass.
0.00359% of the water used for washing is assumed to be lost during the washing
process.
120
CALCULATIONS
Sorting and Sand Removing
Cassava receipt rate = 15000kg/day
๐ด๐๐ข๐๐ก ๐๐ ๐ค๐๐ ๐ก๐ ๐๐๐๐๐ฃ๐๐, ๐ด๐1 =1
100ร ๐ด๐1 = 0.01 ร 15000 = 150๐พ๐/๐๐๐ฆ
๐ ๐๐๐ก๐ ๐ด๐๐ก๐๐ ๐๐๐๐ก๐๐๐, ๐ด๐2 = ๐ถ๐๐ ๐ ๐๐ฃ๐ ๐๐๐๐๐๐ฃ๐๐ โ ๐๐๐ ๐ก๐ ๐๐๐๐๐ฃ๐๐
= 15000 โ 150 = 14850๐พ๐/๐๐๐ฆ
Washing and Peeling
Amount of water used for washing, Bi1 =141.4
100ร ๐ด๐1 = 1.414 ร 14850
= 20997.9๐พ๐/๐ท๐ด๐
๐ด๐๐๐ข๐๐ก ๐๐ ๐ค๐๐ ๐ก๐ ๐ค๐๐ก๐๐, ๐ต๐2 =99.641
100ร ๐ต๐1 = 0.99641 ร
20997.9๐พ๐
๐๐๐ฆ
= 20922.52๐พ๐/๐๐๐ฆ
๐ด๐๐๐ข๐๐ก ๐๐ ๐๐๐๐๐ , ๐ต๐1 =1.5
100ร ๐ด๐2 = 0.015 ร 14850
=222.75๐พ๐
๐๐๐ฆ
๐๐๐๐๐๐ ๐ ๐๐๐ก๐ , ๐ต๐2 = ๐ด02 + ๐ต๐1 โ (๐ต๐1 + ๐ต๐2)
= 14850 + 20997.9 โ (20997.9 + 222.75)
= 14702.63๐พ๐
๐๐๐ฆ
121
RASPING AND EXTRACTION
ASSUMPTIONS
Mass of water added to rasper is 36.5% of the total mass roots fed into the rasper
The mass Sulphur water added to the slurry for extraction is 32% of the total mass of
roots rasped
The coarse portion of the starch slurry removed is 35% of the mass of slurry after
rasping.
122
CALCULATIONS
Root Rasping
๐ค๐๐ก๐๐ ๐๐๐๐๐ ๐ก๐ ๐๐๐ ๐๐๐, ๐ถ๐2 =36.5
100ร ๐ถ๐1
๐๐ข๐ก ๐ถ๐1 = ๐ต๐2 =14702.63๐พ๐
๐๐๐ฆ
๐ถ๐2 = 0.365 ร 14702.63 ๐๐/๐๐๐ฆ
=5366.46๐๐
๐๐๐ฆ
๐ ๐๐ ๐๐๐ ๐๐ข๐๐, ๐ถ๐1 = ๐ถ๐1 + ๐ถ๐2
= 14702.63 + 5366.46 = 20069.09๐พ๐
๐๐๐ฆ
1st Extraction
๐ ๐ข๐๐โ๐ข๐ ๐ค๐๐ก๐๐ ๐๐๐๐๐, ๐ท๐1 =32
100ร ๐ถ๐1 = 0.32 ร
14702.63 ๐พ๐
๐๐๐ฆ
๐ท๐1 =6422.11๐พ๐
๐๐๐ฆ
๐๐๐ก๐๐ ๐ข๐ ๐๐, ๐ท๐2 =40
100ร ๐ถ๐1 = 0.40 ร
14702.63๐พ๐
๐๐๐ฆ
๐ท๐2 =5881.05๐พ๐
๐๐๐ฆ
๐๐๐๐๐ ๐ ๐ ๐ก๐๐๐โ ๐๐๐๐ฆ๐๐๐๐, ๐ท๐1 =35
100ร (๐ถ๐1 + ๐ท๐1 + ๐ท๐2)
= 0.35 ร (14702.63 + 6422.11 + 5881.05)
123
= 0.35 ร27005.79๐พ๐
๐๐๐ฆ
=11330.29๐พ๐
๐๐๐ฆ
๐น๐๐๐ ๐๐๐ข๐๐๐ฆ, ๐ท๐2 = ๐ท๐1 + ๐ท๐2 + ๐ท๐1
= 6422.11 + 5881.05 + 11330.29
= 21041.97๐พ๐
๐๐๐ฆ
2ND EXTRACTION AND PULP PRESSING
ASSUMPTIONS
Mass of coarse cassava starch is equal to the mass of coarse starch discharged from the
first extraction stage.
Mass of water added to the coarse starch for extraction is 45% of the total mass of coarse
starch input into the extractor.
Fine starch slurry extracted is assumed to be 65% of total slurry and water introduced
into the extractor.
124
CALCULATIONS
2ND Extraction
๐พ๐๐๐๐ ๐๐๐๐ ,๐ฏ๐๐ =45
100ร ๐ป๐1
But ๐ป๐1 = ๐ท๐1 = 11330.29๐พ๐
๐๐๐ฆ
Hence ๐ป๐2 = 0.45 ร 11330.29๐พ๐
๐๐๐ฆ
= 5098.63 ๐พ๐/๐๐๐ฆ
๐ญ๐๐๐ ๐๐๐๐๐๐ ๐๐๐๐๐๐๐๐๐ ,๐ป๐1 = 65
100ร ๐ป๐1 = 0.65 ร 11330.29
= 7364.69๐พ๐
๐๐๐ฆ
125
๐ช๐๐๐๐๐๐ ๐ท๐๐๐,๐ป๐2 = ๐ป๐1 + ๐ป๐2 โ ๐ป๐1
= 11330.29 + 5098.63 โ 7364.69 = 9064.23๐พ๐
๐๐๐ฆ
PULP PRESS
๐ ๐ก๐๐๐โ ๐ ๐๐ข๐๐๐ฆ ๐๐๐ก๐๐๐๐๐, ๐ผ๐1 =31
100ร ๐ป๐2
= 0.31 ร 9064.23 = 2809.91๐พ๐
๐๐๐ฆ
๐ ๐ก๐๐๐โ ๐ถ๐๐๐, ๐ผ๐2 = ๐ป๐2 โ ๐ผ๐1
= 9064.23 โ 2809.91 = 6254.32๐พ๐
๐๐๐ฆ
SEPARATION AND DEWATERING
ASSUMPTIONS
The mass of water added to the separator is assumed to be 30% of the mass of fine
slurry extracted.
The mass of Sulphur water added is 15% of the mass of the fine slurry extracted.
The mass of waste water from the separator is assumed to be 67% of the total mass of
feed flows.
The waste water at the dewatering process is 67% of the concentrated slurry.
The input is the fine starch from the 2nd extraction plus the fine starch from the 1st
extraction.
126
CALCULATIONS
Separator
๐ค๐๐ก๐๐ ๐๐๐๐ข๐ก, ๐ธ๐2 =30
100ร (๐ธ๐1 + ๐ป๐1)
= 0.30 ร (21041.97 + 7364.69)
= 4261๐พ๐
๐๐๐ฆ
๐ ๐ข๐๐โ๐ข๐ ๐ค๐๐ก๐๐ ๐๐๐๐ข๐ก, ๐ธ๐3 =15
100ร ๐ธ๐1
= 0.15 ร 28406.65
= 8522๐พ๐
๐๐๐ฆ
127
๐ค๐๐ ๐ก๐๐ค๐๐ก๐๐, ๐ธ๐1 =67
100ร (๐ธ๐1 + ๐ป๐1 + ๐ธ๐2 + ๐ธ๐3)
= 0.67 ร (21041.97 + 7364.69 + 4261 + 8522)
= 27597.07๐พ๐
๐๐๐ฆ
๐๐๐๐๐๐๐ก๐๐๐ก๐๐ ๐ ๐๐ข๐๐๐ฆ, ๐ธ๐2 = ๐ธ๐1 + ๐ป๐1 + ๐ธ๐2 + ๐ธ๐3 โ ๐ธ๐1
= 21041.97 + 7364.69 + 4261 + 8522 โ 27597.07
= 13592.58๐พ๐
๐๐๐ฆ
Dewatering
๐ค๐๐ ๐ก๐๐ค๐๐ก๐๐, ๐น๐1 =67
100ร ๐ธ๐2
= 0.67 ร 13592.58
= 9107.03๐พ๐
๐๐๐ฆ
๐ ๐ก๐๐๐โ ๐๐๐๐, ๐น๐2 = ๐ธ๐2 โ ๐น๐1
= 13592.58 โ 9107.03
= 4485.55๐พ๐
๐๐๐ฆ
This is diluted to produce 6049 kg of starch slurry and sent to the glucose plant.
128
MASS BALANCE FOR THE GLUCOSE PRODUCTION PLANT
LIQUEFACTION PROCESS
ASSUMPTIONS
The starch slurry contains 70% water
Alpha-amylase added is 0.2% of the starch slurry
20% of water is lost during liquefaction
Calcium ions added is small and negligible.
129
CALCULATIONS
LIQUEFACTION
Amount of water loss during liquefaction, wl =20% of Starch slurry, Si
๐ค๐ = 0.2 ร 6048.8 = 1209.8 kg/day
๐๐ + ๐ธ1 = ๐๐1 + ๐ค๐
6048.8 + 12 = ๐๐1 + 1209.8
= 4851 ๐๐/๐๐๐ฆ
partially hydrolyzed starch, ๐๐1 = 4851 ๐๐/๐๐๐ฆ
The quantity of partially hydrolyzed starch from the liquefaction process is 4851 ๐๐/๐๐๐ฆ
130
SACCHARIFICATION
ASSUMPTON
The water loss during the liquefaction process is negligible.
๐01 + ๐ธ2 =, ๐บ01
๐บ๐1 = 4851 + 9.7
๐บ๐1 = 4860.7 ๐๐/๐๐๐ฆ
Crude Glucose, ๐บ๐1 = 4860.7 ๐๐/๐๐
The quantity of Crude Glucose after saccharification process is 4861 kg/day
131
ACIDIFICATION
The amount of Sulphuric Acid used is 5% of the crude glucose =5% x 4861
๐๐ = 243 ๐๐/๐๐๐ฆ
Total mass in =Total mass out
๐บ๐1 + ๐๐ด = ๐บ02
๐บ02 = 4861 + 243
๐บ02 = 50103.7
Acidified Glucose, ๐บ02 = 5103.7 ๐๐/๐๐๐ฆ
The quantity of Glucose after Acidification process is 5104 kg/day
132
ADSORPTION
ASSUMPTION
One per cent of the syrup is adsorbed as residue
The residues,Ro1 is 1% of the incoming glucose syrup=1% ร ๐บ๐2
= 1% ร 5104
= 51๐๐/๐๐๐ฆ
Mass in =Mass out
๐บ02 = ๐บ03 + ๐ 01
133
๐บ๐3 = 5104 โ 51
๐บ๐3 = 5053
Refined Glucose syrup, ๐บ๐3 = 5053 ๐๐/๐๐๐ฆ
The quantity of the refined Glucose syrup after the adsorption is 5053 ๐๐/๐๐๐ฆ
ION EXCHANGE (CATION)
ASSUMPTION
The residues, Ro2 is 0.3% of the incoming glucose syrup=0.3% ร ๐บ๐3
= 0.3% ร 5104
= 15.2 ๐๐/๐๐๐ฆ
134
CALCULATIONS
๐๐๐ ๐ ๐ผ๐ = ๐๐๐ ๐ ๐๐ข๐ก
๐บ๐3 = ๐บ๐4 + ๐ 02
๐บ๐4 = 5053 โ 15.2
Refined Glucose syrup, ๐บ๐4 = 5037.8 ๐๐/๐๐๐ฆ
The quantity of the De-ionised Glucose syrup after the adsorption is 5038 ๐๐/๐๐๐ฆ
ION EXCHANGE (ANION)
ASSUMPTION
The residues,Ro2 is 0.3% of the incoming glucose syrup=0.3% ร ๐บ๐3
= 0.3% ร 5104
= 15.2 ๐๐/๐๐๐ฆ
135
CALCULATIONS
Mass in = Mass out
๐บ๐4 = ๐ 03 + ๐บ๐5
๐บ๐5 = ๐บ๐4 โ ๐ ๐5
๐บ๐5 = 5037.8 โ 15
final glucose, ๐บ๐5 = 5022.8 ๐๐/๐๐๐ฆ
The quantity of the Refined Glucose syrup after the adsorption is 5023 ๐๐/๐๐๐ฆ
136
MASS BALANCE FOR THE SORBITOL PREPARATION PLANT
HYDROGENATION
ASSSUMPTIONS
The ratio of dextrose feed to hydrogen gas is 1:5
The catalyst in the reaction has no effect on material balance.
NB. Hydrogenation reaction is a chemical reaction and so the general assumption of mass
entering equals mass out at steady state is not applicable.
The Dextrose Equivalence (DE) is the measure of the reducing sugars in the sorbitol relative to
the dextrose expressed as a percentage.
CALCULATION
DATA
Molar mass of dextrose syrup (C6H12O6) is 180.1559
Molar mass of sorbitol (C6H14O6) is 182.17
Molar mass of hydrogen gas (H2) is 2.016
137
C6H12O6 + H2 C6H14O6
The theoretical yield of sorbitol can be calculated by knowing whether dextrose or hydrogen is
the limiting reactant. For the above equation, the stoichiometric ratio of reactants
๐ (๐ถ6๐ป12๐6)
๐(๐ป2)= 1
๐ผ๐๐๐ก๐๐๐ ๐๐๐๐ข๐๐ก ๐๐ ๐๐๐ฅ๐ก๐๐๐ ๐ = ๐ท๐๐ฅ๐ก๐๐๐ ๐ ๐๐๐๐ค ๐๐๐ก๐ ร1
๐๐๐๐๐ ๐๐๐ ๐ ๐๐ ๐๐๐ฅ๐ก๐๐๐ ๐
= 5503 ร1
180.1559
= 30.5 ๐๐๐ ๐๐ ๐๐๐ฅ๐ก๐๐๐ ๐ ๐ ๐ฆ๐๐ข๐.
๐ผ๐๐๐ก๐๐๐ ๐๐๐๐ข๐๐ก ๐๐ โ๐ฆ๐๐๐๐๐๐ = ๐ป๐ฆ๐๐๐๐๐๐ ๐๐๐๐ค ๐๐๐ก๐ ร1
๐๐๐๐๐ ๐๐๐ ๐ ๐๐ โ๐ฆ๐๐๐๐๐๐
= 1005 ร1
2.016
= 498.3 ๐๐๐ ๐๐ โ๐ฆ๐๐๐๐๐๐ ๐๐๐
๐ ๐๐ก๐๐ ๐๐ ๐กโ๐ ๐๐๐๐ก๐๐๐ ๐๐๐๐ข๐๐ก๐ ๐๐ ๐๐๐๐๐ก๐๐๐ก = ๐ (๐ถ6๐ป12๐6)
๐(๐ป2)
= 30.5
498.3= 0.061
๐๐๐๐๐ ๐กโ๐ ๐๐๐ก๐๐ ๐๐ ๐กโ๐ ๐๐๐๐ก๐๐๐ ๐๐๐๐ข๐๐ก ๐๐ ๐๐๐๐๐ก๐๐๐ก ๐๐ ๐๐๐ ๐ ๐กโ๐๐ ๐กโ๐ ๐ ๐ก๐๐๐โ๐๐๐๐๐ก๐๐๐ ๐๐๐ก๐๐, ๐๐ก ๐๐๐๐๐๐๐ ๐๐๐ฅ๐ก๐๐๐ ๐ ๐๐
๐กโ๐ ๐๐๐๐๐ก๐๐๐ ๐๐๐๐๐ก๐๐๐ก. (๐โ๐ ๐๐๐๐ข๐๐ก ๐๐ ๐๐๐ฅ๐ก๐๐๐ ๐ ๐๐๐๐๐ก๐๐๐ ๐ค๐๐๐ ๐๐๐ก๐๐๐๐๐๐ ๐กโ๐ ๐๐๐๐ข๐๐ก ๐๐ ๐ ๐๐๐๐๐ก๐๐ )
138
๐โ๐ ๐กโ๐๐๐๐๐ก๐๐๐๐ ๐ฆ๐๐๐๐ ๐๐ ๐ ๐๐๐๐๐ก๐๐ ๐๐ ๐๐๐๐๐ = ๐๐๐๐๐ ๐๐ ๐๐๐ฅ๐ก๐๐๐ ๐ ร1 ๐๐๐ ๐๐ ๐ ๐๐๐๐๐ก๐๐
1 ๐๐๐ ๐๐ ๐๐๐ฅ๐ก๐๐๐ ๐
= 30.5 ๐๐๐ ๐๐ ๐ ๐๐๐๐๐ก๐๐.
๐โ๐ ๐๐๐ ๐ ๐๐ ๐ ๐๐๐๐๐ก๐๐ ๐๐๐๐๐ข๐๐๐ = ๐๐๐๐๐ ร ๐๐๐๐๐ ๐๐๐ ๐
= 30.5 ร 182.17
๐๐๐ ๐ ๐๐ ๐ ๐๐๐๐๐ก๐๐ ๐๐๐๐๐ข๐๐๐ ๐๐ ๐ ๐๐๐ฆ ๐๐ 5474๐๐
๐๐๐ฆ
139
EVAPORATION
ASSUMPTIONS
Initial sorbitol content before evaporation (S) is 55%.
Sorbitol content (S) increased to 70% after evaporation.
All excess hydrogen injected into the system is expelled from the product by reducing
pressure.
CALCULATIONS
Total mass balance on evaporation
Total mass in = Total mass out
๐ต = ๐ถ + ๐ท
B =Flow rate of sorbitol solution =5474 kg/day
C = Flow rate of water evaporation
D = Flow rate of concentrated sorbitol solution
5474๐๐
๐๐๐ฆ= ๐ถ + ๐ท โฆ. (1)
140
Component mass balance on evaporation
Sorbitol content (S)
Component mass in = Component mass out
๐๐๐ ๐ ๐๐ ๐ ๐๐๐๐๐ก๐๐ = % ๐ ๐๐๐๐๐ก๐๐ ร ๐๐๐๐ค ๐๐๐ก๐
๐ต(๐๐ต) = ๐ถ(๐๐ถ) + ๐ท(๐๐ท)
๐๐ต = % ๐๐๐๐๐๐ก๐๐ ๐๐ ๐ ๐๐๐๐๐ก๐๐ ๐ ๐๐๐ข๐ก๐๐๐ = 55%
๐๐ท = % ๐๐๐๐๐๐ก๐๐ ๐๐ ๐๐๐๐๐๐๐ก๐๐๐ก๐๐ ๐ ๐๐๐๐๐ก๐๐ = 70%
5474(0.55) = ๐ท(0.70)
๐ต๐ข๐ก ๐๐๐๐ (1), ๐ท =3010.7
0.7
๐ท = 4301๐๐
๐๐๐ฆ
๐ถ = 1173๐๐
๐๐๐ฆ
141
CRYSTALLISATION
ASSUMPTIONS
Moisture content (M) is reduced from 30% to 10% after crystallization.
The product of the crystallization is a mixture of sorbitol crystals in the mother liquor.
CALCULATIONS
Mass balance on crystallization
Total mass balance on crystallization
Total Mass in = Total Mass out
๐ท = ๐ธ + ๐น . . (1)
D = Flow rate of concentrated sorbitol = 4301 kg/day
E= Flow rate of water removed during the process
F = Flow rate of sorbitol crystals produced.
142
Component Mass balance
Component mass in = Component mass out
Sorbitol Content (M)
๐๐๐ ๐ ๐๐ ๐ ๐๐๐๐๐ก๐๐ = % ๐ ๐๐๐๐๐ก๐๐ ร ๐น๐๐๐ค ๐๐๐ก๐
๐ท(๐๐ท) = ๐ธ(๐๐ธ) + ๐น(๐๐น)
๐๐ท = % ๐ ๐๐๐๐๐ก๐๐ ๐๐๐๐ก๐๐๐ก ๐๐ ๐๐๐๐๐๐๐ก๐๐๐ก๐๐ ๐ ๐๐๐๐๐ก๐๐ = 70%
๐๐น = % ๐ ๐๐๐๐๐ก๐๐ ๐๐ ๐๐๐ฆ๐ ๐ก๐๐ ๐ ๐๐๐๐๐ก๐๐ = 90%
4301(0.7) = 0.90๐น
3010.7 = 0.90๐น
๐น = 3345 ๐๐
๐๐๐ฆ
๐น =956 ๐๐
๐๐๐ฆ
143
FILTRATION
ASSUMPTIONS
Filtration only retains sorbitol crystals leaving the mother liquor to pass.
The filtered crystals have a moisture content of 5%
CALACULATIONS
Mass balance on filtration
Total Mass balance
Total Mass in = Total Mass out
๐น = ๐บ + ๐ป โฆโฆโฆ . . (1)
F = Flow rate of crystal sorbitol = 3345kg/day.
G = Flow rate of filtered liquid.
H = Flow rate of pure Sorbitol crystals.
Component mass balance
144
Component mass in = component mass out
Moisture content (M)
๐๐๐ ๐ ๐๐ ๐๐๐๐๐๐ก๐๐ = % ๐ ๐๐๐๐๐ก๐๐ ร ๐น๐๐๐ค ๐๐๐ก๐
๐น(๐๐น) = ๐บ(๐๐บ) + ๐ป(๐๐ป)
๐๐น = % ๐ ๐๐๐๐๐ก๐๐ ๐๐ ๐๐๐ฆ๐ ๐ก๐๐ ๐ ๐๐๐๐๐ก๐๐ = 90%
๐๐น๐ป = % ๐ ๐๐๐๐๐ก๐๐ ๐๐ ๐ ๐๐๐๐๐ก๐๐ ๐๐๐๐ = 95%
3345(0.90) = 0.95 ๐ป
๐ป =3169 ๐๐
๐๐๐ฆ
๐บ =176 ๐๐
๐๐๐ฆ
145
DRYING
ASSUMPTION
Moisture content of sorbitol must not be more than 1% and so the main assumption is that
the final product has a moisture content of 1%.
CALCULATIONS
Total mass balance on drying
๐๐๐ก๐๐ ๐๐๐ ๐ ๐๐ = ๐ก๐๐ก๐๐ ๐๐๐ ๐ ๐๐ข๐ก
๐ป = ๐ผ + ๐ฝ
๐โ๐๐๐, ๐ป = ๐กโ๐ ๐๐๐ ๐ ๐๐๐๐ค ๐๐๐ก๐ ๐๐ ๐๐๐๐ ๐ก ๐ ๐๐๐๐๐ก๐๐ ๐๐๐๐๐
๐ผ = ๐น๐๐๐ค ๐๐๐ก๐ ๐๐ ๐๐๐๐ ๐ก๐ข๐๐ ๐๐๐๐๐ฃ๐๐ ๐๐ข๐๐๐๐ ๐๐๐ฆ๐๐๐
๐ฝ = ๐น๐๐๐ค ๐๐๐ก๐ ๐๐ ๐๐๐๐๐ ๐ ๐๐๐๐๐ก๐๐ ๐๐๐ฆ๐ ๐ก๐๐๐
Component mass balance
Component mass in = component mass out
146
Moisture content (M)
๐๐๐ ๐ ๐๐ ๐ ๐๐๐๐๐ก๐๐ = % ๐ ๐๐๐๐๐ก๐๐ ร ๐๐๐๐ค ๐๐๐ก๐
๐๐ป = % ๐ ๐๐๐๐๐ก๐๐ ๐๐ ๐ ๐๐๐๐๐ก๐๐ ๐๐๐๐ = 95%
๐๐ฝ = % ๐ ๐๐๐๐๐ก๐๐ ๐๐ ๐๐๐๐๐ ๐ ๐๐๐๐๐ก๐๐ ๐๐๐ฆ๐ ๐ก๐๐ = 99%
๐ป(๐๐ป) = ๐ผ(๐๐ผ) + ๐ฝ(๐๐ฝ)
3169(0.95) = 0.99๐ฝ
๐ฝ = 3169(0.95)
0.99 =
3041 ๐๐
๐๐๐ฆ
๐ผ = 125 ๐๐
๐๐๐ฆ
147
MILLING
ASSUMPTION
1.0% of the crystals are assumed to go waste during the milling process.
CALCULATIONS
Total mass balance on milling
๐๐๐ก๐๐ ๐๐๐ ๐ ๐๐ = ๐ก๐๐ก๐๐ ๐๐๐ ๐ ๐๐ข๐ก
๐ฝ = ๐พ + ๐ฟ
Where J = the flow rate of the dried sorbitol crystal = 3041kg/day
K = the flow rate of waste sorbitol
L = the flow rate of the fine sorbitol powder
๐ด๐๐๐ข๐๐ก ๐๐ ๐ค๐๐ ๐ก๐ (๐พ) = 1.0
100 ร
3041๐๐
๐๐๐ฆ
๐ด๐๐๐ข๐๐ก ๐๐ ๐ค๐๐ ๐ก๐ (๐พ) โ 30.41 ๐๐
๐๐๐ฆ
148
๐น๐๐๐ ๐กโ๐ ๐ก๐๐ก๐๐ ๐๐๐ ๐ ๐๐๐๐๐๐๐, ๐ฟ =3041๐๐
๐๐๐ฆ โ
30.41 ๐๐
๐๐๐ฆ
๐โ๐ ๐๐๐๐ค ๐๐๐ก๐ ๐๐ ๐๐๐๐ ๐ ๐๐๐๐๐ก๐๐ ๐๐๐ค๐๐๐ (๐ฟ) โ3000 ๐๐
๐๐๐ฆ
PACKAGING
ASSUMPTION
No mass loss or gain during packaging.
CALCULATION
Output of the plant = 3000kg of sorbitol daily.
Quantity in 50kg bags = 3000 ๐๐
50 ๐๐
๐๐ข๐๐๐ก๐๐ก๐ฆ ๐๐ ๐ ๐๐๐๐๐ก๐๐ = 60 ๐๐๐๐
๐๐๐ฆ
๐โ๐๐๐๐๐๐๐ ๐๐๐ ๐ ๐๐๐๐ ๐๐5023๐๐
๐๐๐ฆ, 60 ๐๐๐๐ ๐๐ ๐ 50๐๐ ๐ ๐๐๐๐๐ก๐๐ ๐๐๐ฆ๐ ๐ก๐๐๐ ๐ค๐๐๐ ๐๐ ๐๐๐ก๐๐๐๐๐.
149
ENERGY BALANCE
ENERGY BALANCE FOR THE DEXTROSE PRODUCTION PLAN
LIQUEFACTION PROCESS
The energy balance on the various stages and equipment is based on the law of conservation of
energy.
โ๐ธ = ๏ฟฝ๏ฟฝ(โ๐ป + โ๐พ๐ธ + โ๐๐ธ) ยฑ ๐ ยฑ ๐
At steady state, โ๐ธ = 0
There is no movement of equipment, velocity is zero hence โ๐พ๐ธ = 0
โ๐๐ธ = 0
The above equation reduces to
๐ = ๏ฟฝ๏ฟฝโ๐ป + ๐
๏ฟฝ๏ฟฝโ๐ป = ๐น๐ ๐ถ๐๐ โ๐
(Siebel 1892), proposed the following formula for estimating the specific heat values above and
below freezing;
๐ถ๐ = 3.35๐๐ค + 0.84 (๐พ๐ฝ/๐๐๐พ)โฆโฆโฆโฆโฆโฆโฆโฆโฆ( ๐๐๐๐ฃ๐ ๐๐๐๐๐ง๐๐๐)
๐ถ๐ = 1.26๐๐ค + 0.84 (๐พ๐ฝ
๐๐๐พ)โฆโฆโฆโฆโฆโฆโฆโฆโฆ( ๐๐๐๐๐ค ๐๐๐๐๐ง๐๐๐)
Where ๐๐ค = ๐กโ๐ ๐๐๐ ๐ ๐๐๐ก๐๐ ๐๐ ๐ค๐๐ก๐๐ ๐๐ ๐กโ๐ ๐๐๐๐๐ข๐๐ก
150
Using equation (1)
The specific capacity of the starch slurry at 70% of moisture content; that is ๐๐ค = 0.7
๐ถ๐๐ = 3.35(0.7) + 0.84 (๐พ๐ฝ
๐พ๐๐พ)
๐ถ๐๐ = 3.185 (๐พ๐ฝ
๐๐๐พ)
๐ถ๐๐ ๐๐ ๐กโ๐ ๐ ๐๐๐๐๐๐๐ โ๐๐๐ก ๐๐๐๐๐๐๐ก๐ฆ ๐๐ ๐กโ๐ ๐ ๐ก๐๐๐โ ๐ ๐๐ข๐๐๐ฆ
๏ฟฝ๏ฟฝโ๐ป = ๐น๐ ๐ถ๐๐ โ๐
๏ฟฝ๏ฟฝโ๐ป = 6048.8 ร 3.185 ร (97 โ 25)
= 1483437๐พ๐ฝ
๐๐๐ฆ(
1๐๐๐ฆ ๐ฅ 1โ๐
24 โ๐ ร 3600๐ )
= 17.2 ๐๐
Therefore, the quantity of energy supplied by the steam is 17.2 kW
The quantity of steam required to supply this energy
17.2 =ms ฮปs
๐๐ = 11.5
๐๐ฌ
151
๐๐๐๐ ๐ ๐ก๐๐๐ ๐ก๐๐๐๐, ๐ ๐ก๐๐๐ ๐๐ก 120โ โ๐๐ ๐๐ฌ = ๐๐๐๐
=17.2
2258= 7.6 ร 10โ3 kg/s
= (7.6 ร 10โ3) ๐๐
๐ (3600๐ )
1โ๐
๐๐ = 27.4 ๐๐/โ๐
Therefore18.4 kg/hr of steam is needed to supply this energy
SURFACE AREA REQUIRED
๐ = ๐๐ด๐ โ๐
๐ = 17.2 ๐๐
๐ = ๐๐ฃ๐๐๐๐๐ โ๐๐๐ก ๐ก๐๐๐๐ ๐๐๐ ๐๐๐๐๐๐๐๐๐๐๐ก = 300 ๐/๐2โ
Mean liquid temperature ๐๐ =๐1+๐2
2=
97+25
2
= 61โ
152
Steam at 100 kpa has temperature, ๐๐ = 120โ
โ๐ = Mean temperature difference = ๐๐ โ ๐๐
= 120 โ 61
โ๐ = 59โ
๐ด๐ =๐
โ๐๐
=17.2 ร103
59 ร300
๐ด๐ = 0.97 ๐2
The work done by the stirrer, W
๐ท
๐= 0.4
T=2481mm
D=0.4 x 2481= 992.4mm
Where D is the diameter of the agitator and T is the diameter of the vessel
Agitator Speed (N) = 53.20 rpm
Power Number(๐๐) = 1.370
Power(P) = Np ร ฯ ร ๐3 ร ๐ท5
Power(P) = 1.370 ร 1200 ร 53.203 ร 992.45
153
Power(P) = 1.10 kW
Assuming loading of 80% efficiency, motor power required
Motor Power = 1.10/ 0.8
= 1.38 kW
Hence
๐ = ๏ฟฝ๏ฟฝโ๐ป + ๐
๐ = 17.2 + 1.38
๐ = 18.6 ๐๐
The total amount of energy needed during liquefaction is 18.6 ๐๐
SACCHARIFICATION PROCESS
During this process, the partially hydrolysed starch at 97ยฐ๐ถ is reduced to 60ยฐ๐ถ, thus, there is heat
energy loss.
๐ถ๐ = 3.35๐๐ค + 0.84 (๐พ๐ฝ/๐๐๐พ)โฆโฆโฆโฆโฆโฆโฆโฆโฆ( ๐๐๐๐ฃ๐ ๐๐๐๐๐ง๐๐๐)
๐ถ๐๐๐1= 3.35 ร 0.50 + 0.84 (๐พ๐ฝ/๐๐๐พ)
๐ถ๐๐๐1=2.52 (๐พ๐ฝ/๐๐ยฐ๐ถ)
The specific heat capacity of the partially hydrolysed starch, ๐ถ๐๐๐1 is 2.68 (๐พ๐ฝ/๐๐ยฐ๐ถ)
154
The flow rate of partially hydrolysed starch, ๐๐1 = 4851
๏ฟฝ๏ฟฝโ๐ป = ๐๐1 ร ๐ถ๐๐๐1ร โ๐
= 4851 ร 2.52 ร (60 โ 97)
= โ452307๐พ๐ฝ
๐๐๐ฆ(
1๐๐๐ฆ ๐ฅ1โ๐
24โ๐ ๐ฅ 3600๐ )
= โ5.2 ๐๐
Therefore 5.2 kW of heat energy is removed.
The quantity of water required to remove this amount of energy. The water enters at a
temperature of 28ยฐ๐ถ and leaves at 38ยฐ๐ถ
๐ = ๐๐ค๐ถ๐๐คโ๐
๐๐ค =๐
๐ถ๐๐คโ๐
=5.2
4.187(38โ28)
=0.12 kg/s
= 430 ๐๐/โ๐
Therefore, 430 ๐๐/โ๐ of water at 28ยฐ๐ถ is needed.
155
AREA OF STEAM JACKET REQUIRED
๐ = ๐๐ด๐ โ๐
๐ = 5.2 ๐๐
๐ = ๐๐ฃ๐๐๐๐๐ โ๐๐๐ก ๐ก๐๐๐๐ ๐๐๐ ๐๐๐๐๐๐๐๐๐๐๐ก = 300 ๐/๐2โ
Mean liquid temperature ๐๐ =๐1+๐2
2=
97+60
2
= 79.5โ
Steam at 100 kpa has temperature, ๐๐ = 120โ
โ๐ = Mean temperature difference = ๐๐ โ ๐๐
= 120 โ 79.5
โ๐ = 41.5โ
๐ด๐ =๐
โ๐๐
=5.2 ร 103
41.5 ร 300
๐ด๐ = 0.4 ๐2
The work done by the propeller, W
๐ท
๐= 0.4
T=2167.7mm
D=0.4 x 2167.7= 867.08 mm
156
Where D is the diameter of the agitator and T is the diameter of the vessel
Agitator Speed (N) = 60.90 rpm
Power Number(๐๐) = 1.370
Power(P) = Np ร ฯ ร ๐3 ร ๐ท5
Power(P) = 1.370 ร 1540 ร 60.903 ร 867.085
Power(P) = 1.08 kW
Assuming loading of 80% efficiency, motor power required
Motor Power = 1.08
= 1.35 kW
Hence
๐ = ๏ฟฝ๏ฟฝโ๐ป + ๐
๐ = โ5.2 + 1.35
๐ = โ3.85 ๐๐
There are -3.85๐๐ of energy lost during saccharification.
157
PLATE HEAT EXCHANGER WITH FOUR PLATES
๐๐ค = 0.5
๐ = 3.35๐๐ค + 0.84 (๐พ๐ฝ/๐๐๐พ)โฆโฆโฆโฆโฆโฆโฆโฆโฆ( ๐๐๐๐ฃ๐ ๐๐๐๐๐ง๐๐๐)
๐ถ๐๐บ02= 3.35 ร 0.5 + 0.84 (๐พ๐ฝ/๐๐๐พ)
๐ถ๐๐บ02=2.52 (๐พ๐ฝ/๐๐ยฐ๐ถ)
The specific heat capacity of the partially hydrolysed starch, ๐ถ๐๐๐1 is 2.52 (๐พ๐ฝ/๐๐ยฐ๐ถ)
๐ = ๐บ๐2 ร ๐ถ๐๐บ01ร โ๐
= ๐บ๐2 ร ๐ถ๐๐บ01ร (๐2 โ ๐1)
๐ = 5103.9 ร 2.52 ร (77 โ 60)
= 218651๐พ๐ฝ
๐๐๐ฆ(
1๐๐๐ฆ ร 1โ๐
24 ๐๐๐ฆร3600๐ )
= 2.5 ๐๐
158
The water enters at a temperature of ๐๐ค1 = 110ยฐ๐ถ and leaves at ๐๐ค2 =95ยฐ๐ถ
๐ = ๐๐ค๐ถ๐๐คโ๐
๐๐ค =๐
๐ถ๐๐ค(๐๐ค1โ๐๐ค2 )
=2.5
4.187(110โ95)
=0.04 kg/s
= 145 ๐๐/โ๐
Therefore, 89.7 ๐๐/โ๐ of water at ยฐ๐ถ is needed.
โ๐๐ =(๐๐ค2 โ๐1)โ(๐๐ค1 โ๐2)
ln [(๐๐ค2 โ๐1)
(๐๐ค1 โ๐2)]
=(95โ60)โ(110โ77)
ln [(95โ601)
(110โ77)]
= 34ยฐ๐ถ
159
Calculating the Area of the heat exchanger
๐ = ๐๐ด๐ โ๐๐
The Overall heat transfer coefficient ๐ = 4.23 ๐๐/๐2๐พ
๐ด๐ =2.5
4.23ร34
= 0.02 ๐2
Therefore, the area of the heat exchanger is 0.02 ๐2
160
ENERGY BALANCE FOR THE SORBITOL PREPARATION PLANT
ENERGY BALANCE ON THE SHELL AND TUBE HEAT EXCHANGER.
The glucose syrup mixed with hydrogen is passed through a shell and tube heat exchanger to
increase the temperature to 100 โ. this will ensure that the time spent in the reactor is reduced to
prevent excessive temperature causing caramelisation.
The energy balance on the equipment is based on the law of conservation of energy.
โ๐ธ = ๏ฟฝ๏ฟฝ(โ๐ป + โ๐พ๐ธ + โ๐๐ธ) ยฑ ๐ ยฑ ๐
At steady state, โ๐ธ = 0
There is no movement of equipment, โ๐พ๐ธ = 0
Since there are no elevations, โ๐๐ธ = 0
Since there is no work done, ๐ = 0
The above equation reduces to
๐ = ๐โ๐ป
161
DATA
Temperature inlet, T1 = 70 โ
Temperature outlet T2 = 100โ
Temperature of steam Ts = 110 โ
๐๐๐๐ ๐ = ๐โ๐ป
๐ = 5023 ร 2.68 ร (100 โ 70)
๐ = 403849 ๐๐ฝ
This means the quantity of heat needed by the system is 403849 ๐๐ฑ
๐๐ = 403849 ๐๐ฝ
24 ร 3600 ๐ = 4.67 ๐๐ฝ
The mean heat transfer rate, ๐๐บ = 4.67 kW
โ๐๐ =(๐๐ โ ๐1) โ (๐๐ โ ๐2)
ln [(๐๐ โ ๐1)(๐๐ โ ๐2)
]
โ๐๐ =(110 โ 70) โ (110 โ 100)
ln [(110 โ 70)(110 โ 100)
]
โ๐๐ = 21.6 โ
162
Using the equation ๐ = ๐ด๐โ๐๐
Taking the overall heat transfer coefficient U = 0.50 kW /m2K
๐ด =๐
๐โ๐๐=
4.67
0.5 ร 21.6
๐ด = ๐. ๐๐ m2
Mass of steam
๐ = ๐
๐ถ๐โ๐=
4.67
2.68 ร (100 โ 70)
๐ = 222.4 ๐๐/โ
The mass of steam needed to raise the temperature from 70 โ to 100โ using th shell and tube
heat exchanger is ๐๐๐. ๐ ๐๐/๐
HYDROGENATION
The energy balance on the equipment is based on the law of conservation of energy.
โ๐ธ = ๏ฟฝ๏ฟฝ(โ๐ป + โ๐พ๐ธ + โ๐๐ธ) ยฑ ๐ ยฑ ๐
At steady state, โ๐ธ = 0
There is no movement of equipment, โ๐พ๐ธ = 0
Since there are no elevations, โ๐๐ธ = 0
The above equation reduces to
๐ = ๐โ๐ป + ๐
163
FORMS OF HEAT IN THE REACTOR
In the hydrogenation process, there are two types of heat (Q) generated in the reactor. The heat
generated by the system (QG) and the heat added to the system in the form of steam(QS).
The total energy ๐ = ๐๐บ+๐๐
DATA FOR CALCULATING THE HEAT ADDED TO THE SYSTEM (QS)
Mass flow rate of sorbitol, m =5474kg/day
Inlet temperature, T1 = 100 โ
Outlet temperature T2 = 130 โ
Specific heat capacity of sorbitol using Seibel, ๐ถ๐ = 0.875
Enthalpy of inlet steam at 100 โ , H1 =2676kJ/kg
Enthalpy of outlet steam at 130โ, H2 = 2787 kJ/kg
๐๐ = ๐โ๐ป + ๐
๐โ๐ป = 5474 ร (2787 โ 2676) + ๐
๐ธ๐บ = ๐๐๐๐๐๐ + ๐พ
164
THE WORK DONE BY THE STIRRER, W
The ratios of the diameter of the stirrer (๐ท๐) to the diameter of the tank (๐ท๐) range from 0.3 to
12. Taking a ratio of 0.5. But the tank has a diameter ๐ท๐ of 1.8m
๐ท๐
๐ท๐= 0.5
๐ท๐ = 0.5 ร 1.8 = 0.90 ๐
DATA
Density of sorbitol, = 1490 kg/m3
Agitator Speed (N) = 8.4๐๐ โ1 = 180 rpm
Power Number(๐๐) = 1.370
Using the equation, Power(P) = Np ร ฯ ร ๐3 ร ๐ท5
Power(P) = 0.61 kW
Assuming 85% efficiency, the motor power required is 0.61
0.85 = 0.72 kW
Therefore the work done by the stirrer,๐ = 0.72 ร 24 ร 3600 = 62208 kJ
Substituting, W= 62208 kJ
๐๐ = 607614 + ๐
๐๐ = 607614 + 62208
๐ธ๐บ = ๐๐๐๐๐๐ kJ.
165
This means the quantity of heat needed by the system is ๐๐๐๐๐๐ ๐คJ.
๐๐ = 669822 ๐๐ฝ
24 ร 3600 ๐
๐๐บ = ๐. ๐๐ ๐๐พ
The mean heat transfer rate, ๐๐บ = ๐. ๐๐ kW
The quantity of steam needed to provide this energy ๐๐
๐๐ =๏ฟฝ๏ฟฝโ๐ป
ฮปs ๐คโ๐๐๐ ฮปs at 130 โ = 2787kJ/kg
๐๐ =7.75
2787 = 9.08๐๐/โ
Therefore, the amount of steam required to supply the energy is = 9.08 ๐๐/๐
HEAT GENERATED BY THE SYSTEM (QG)
๐ถ6๐ป12๐6 + ๐ป2 โ โ โ ๐ถ6๐ป14๐6
๐ธ๐๐กโ๐๐๐๐ฆ ๐๐๐๐๐๐๐๐ก๐๐๐ ๐๐ ๐๐๐ข๐๐๐ ๐, โ๐ป๐ถ (๐ถ6๐ป12๐6) = 2805 ๐๐ฝ๐๐๐โ1
๐ธ๐๐กโ๐๐๐๐ฆ ๐๐ ๐๐๐๐๐๐ก๐๐๐ ๐๐ โ๐ฆ๐๐๐๐๐๐, โ๐ป๐ถ (๐ป2) = โ285.5 ๐๐ฝ๐๐๐โ1
๐ธ๐๐กโ๐๐๐๐ฆ ๐๐ ๐๐๐๐๐๐ก๐๐๐ ๐๐ ๐ ๐๐๐๐๐ก๐๐, โ๐ป๐ (๐ถ6๐ป14๐6) = โ3009 ๐๐ฝ๐๐๐โ1
๐๐๐๐ ๐ฏ๐๐๐ ๐๐๐,
โ๐ป๐๐ฅ๐ = โโ๐ป๐๐๐๐๐ข๐๐ก โ โโ๐ป๐๐๐๐๐ก๐๐๐ก
Where โโ๐ป๐๐๐๐๐ข๐๐ก = ๐๐๐กโ๐๐๐๐ฆ ๐๐ ๐ ๐๐๐๐๐ก๐๐ ๐๐๐๐๐๐
166
โโ๐ป๐๐๐๐๐ก๐๐๐ก๐ = ๐๐๐กโ๐๐๐๐ฆ ๐๐ ๐๐๐ข๐๐๐ ๐ + ๐๐๐กโ๐๐๐๐ฆ ๐๐ โ๐ฆ๐๐๐๐๐๐ ๐๐๐๐๐ก๐๐
โ๐ป๐๐ฅ๐ = โ3009 โ (2805 โ 285.5)
โ๐ป๐๐ฅ๐ = โ489.9 ๐๐ฝ๐๐๐โ1
๐ต๐ข๐ก ๐กโ๐ ๐๐๐๐๐ , ๐ ๐๐ ๐ ๐๐๐๐๐ก๐๐ = 30.05 ๐๐๐ ๐๐๐ ๐กโ๐ ๐๐๐๐๐ ๐๐๐ ๐ ,๐ = 182.17
๐ถ๐๐๐ฃ๐๐๐ก๐๐๐ ๐กโ๐ โ๐ป๐๐ฅ๐๐๐๐๐ ๐๐ฝ๐๐๐โ1 ๐ก๐ ๐๐ฝ๐๐โ1 = โ2.73 ๐๐ฝ/๐๐
๐๐บ = ๐ โ๐ป
๐๐บ = 5474(โ2.73)
๐ธ๐ฎ = โ๐๐๐๐๐. ๐๐ ๐๐ฑ
The quantity of energy generated by the system is = โ14941.95 ๐๐ฝ.
๐๐บ =โ14941.95
24 ร 3600= โ๐. ๐๐๐๐๐พ
Substituting the values of ๐ธ๐บ ๐๐๐ ๐ธ๐ฎ
๐ = ๐๐บ+๐๐
๐ธ = 669822 โ 14941.95 = ๐๐๐๐๐๐ ๐๐ฑ
The total amount of energy in the hydrogenation reactor is ๐๐๐๐๐๐ ๐๐ฑ
๐ =๐๐๐๐๐๐ ๐๐ฝ
24ร3600= 7.57 Kw
The mean heat transfer rate, q = = 7.57 kW.
167
Amount of Water Needed to Cool the Reactor
Quantity of water generated by the system QS equals -12247.5 kJ
Assuming cold water enters the vessel at 25โ and leave at 53โ
๐ = ๐๐ถ๐โ๐
Where the heat capacity of water ๐ถ๐ = 4.18๐๐ฝ๐๐โ1
๐ =โ๐๐๐๐๐.๐๐
4.18ร(25โ53)= 4.79 ๐๐/h
The mass of water needed to cool the vessel is 4.79 kg every hour.
ENERGY BALANCE ON EVAPORATION
Evaporation is at a vacuum of 75mmHg
๐๐ข๐ก ๐๐ก 1 ๐๐ก๐ = 760๐๐๐ป๐
๐กโ๐๐๐๐๐๐๐ 75๐๐๐ป๐ =75
760ร 1๐๐ก๐
= 0.0987 ๐๐ก๐
๐๐๐๐ฃ๐๐๐ก๐๐๐ ๐กโ๐๐ ๐ก๐ ๐๐๐ ๐ข๐ ๐๐๐ 1๐๐ก๐ = 101.325, ๐๐๐๐ ๐ ๐ข๐๐ ๐๐ 10.00 ๐๐๐
๐น๐๐๐ ๐ ๐ก๐๐๐ ๐ก๐๐๐๐, ๐ ๐๐๐๐ ๐ ๐ข๐๐ ๐๐ 10.00 ๐๐๐ ๐๐๐๐๐๐ ๐๐๐๐๐ ๐ก๐ ๐ ๐ก๐๐๐๐๐๐๐ก๐ข๐๐ ๐๐ 46โ
168
DATA
Initial temperature of sorbitol, T1= 30 โ
Final temperature of sorbitol, T2 = 46 โ
Cp of sorbitol using the Siebel equation =0.875 kJ/KgK
๐๐ ๐๐๐ ๐ = ๐๐ถ๐โ๐
๐ = 4563.36๐๐ ร 0.875 ร (46 โ 30)
๐ธ = ๐๐๐๐๐. ๐๐๐๐ฑ
The quantity of heat energy needed is 63887.04๐๐ฝ
๐ =63887.04
24 ร 3600= ๐. ๐๐๐ ๐๐พ
Therefore, the heat transfer rate is 0.74kW
Quantity of steam needed.
๐๐ =๐
โ๐
Where he is the latent heat of evaporation at 45C = 2393 kJ/kg
๐๐ =0.74
239= 3.089 ร 104
๐๐ = 1.11๐๐/โ
๐กโ๐ ๐๐๐ ๐ ๐๐ ๐ ๐ก๐๐๐ ๐๐๐๐ข๐๐๐๐ ๐ก๐ ๐๐๐๐ฃ๐๐๐ ๐กโ๐ ๐๐๐๐๐๐ฆ ๐๐ ๐. ๐๐ ๐๐/๐
169
ENERGY BALANCE ON CRYSTALLISATION
The energy balance on the various stages and equipment is based on the law of conservation of
energy.
โ๐ธ = ๏ฟฝ๏ฟฝ(โ๐ป + โ๐พ๐ธ + โ๐๐ธ) ยฑ ๐ ยฑ ๐
At steady state, โ๐ธ = 0
There is no movement of equipment, โ๐พ๐ธ = 0
Since there are no elevations, โ๐๐ธ = 0
Work done W =0
The above equation reduces to
๐ = ๐โ๐ป
DATA
Initial temperature of sorbitol, T1 = 40 โ
Final temperature of sorbitol, T2 = 95โ
Mass flow rate of sorbitol through the crystallizer, m = 4301 kg/day
๐ = 4301 ร 0.875 ร (95 โ 40)
๐ = 206986.6
๐ =206986.6
24 ร 3600= 2.40๐๐
170
MASS OF STEAM REQUIRED TO PRODUCE THIS ENERGY
๐๐ =๐
โ๐ ๐คโ๐๐๐ โ๐ ๐กโ๐ ๐๐๐ก๐๐๐ก โ๐๐๐ก ๐๐ ๐๐๐ฆ๐ ๐ก๐๐๐๐๐ ๐๐ก๐๐๐ ๐๐ก 90 โ = 2270
๐๐ฝ
๐๐
๐๐ =2.40
2270= 0.10
๐๐
โ
The amount of steam needed is 0.10 kg/h
DRYING ENERGY BALANCE
Energy balance of a cabinet dryer, suggest that the thermal energy input to the dryer Q is used to
heat the solid material (Qsh) and the fresh air (Qah). This helps to evaporate moisture Qwe (Z. B.
Maroulis and G. D. Saravacos (2003)
CALCULATING EQUATIONS
๐ = ๐๐ค๐ + ๐๐ โ + ๐๐โ
๐๐ค๐ = ๐น(๐๐ โ ๐)[โ๐ป๐ โ (๐ถ๐๐ฟ โ ๐ถ๐๐)๐]
๐๐ โ = ๐น(๐ถ๐๐ + ๐๐๐ถ๐๐ฟ)(๐ โ ๐๐)
๐๐โ = ๐น๐[๐ถ๐๐ด + ๐๐๐ถ๐๐](๐ โ ๐๐), where
๐น๐ = ๐น(๐๐ โ ๐)
๐ โ ๐๐
171
Where:
F= flow rate of sorbitol crystals (kg/h) =3169
๐๐๐ฆ= 132.04 ๐๐/โ
๐น๐ (๐๐
โ)= flow rate of fresh air entering dryer= 12 kg/h
๐๐=๐๐ค = % moisture content of sorbitol crystals = 0.05
X= % moisture content of dried sorbitol crystals= 0.01
๐๐= ambient humidity= 0.01
Y= drying air humidity= 0.45 kg/kg dry basis
โ๐ป๐= Latent heat at 0oC= 2500 kJ/kg
๐ถ๐๐ฟ= specific heat capacity of water= 4.2 kJ/kg k
๐ถ๐๐= specific heat capacity of vapor= 1.90 kJ/kg k
๐ถ๐๐= specific heat capacity of sorbitol=0.875๐๐ฝ/๐๐ ๐
T= drying air temperature= 65oC
To= ambient temperature= 25oC
172
Calculations
F= 3169 ๐๐
๐๐๐ฆร
1 ๐๐๐ฆ
24 โ๐๐ = 132.04๐๐/โ๐
๐ถ๐๐ ๐ข๐ ๐๐๐ ๐กโ๐ ๐๐๐๐๐ข๐๐๐ ๐๐๐๐๐๐ ๐๐ ๐๐ฆ Siebel (1892) ๐๐๐ estimating the specific heat values
above and below freezing;
๐ถ๐ = 3.35๐๐ค + 0.84 (๐พ๐ฝ/๐๐๐พ)โฆโฆโฆโฆโฆโฆโฆโฆโฆ( ๐๐๐๐ฃ๐ ๐๐๐๐๐ง๐๐๐)
๐ถ๐ = 1.26๐๐ค + 0.84 (๐พ๐ฝ
๐๐๐พ)โฆโฆโฆโฆโฆโฆโฆโฆโฆ( ๐๐๐๐๐ค ๐๐๐๐๐ง๐๐๐)
Where ๐๐ค = ๐กโ๐ ๐๐๐ ๐ ๐๐๐ก๐๐ ๐๐ ๐ค๐๐ก๐๐ ๐๐ ๐กโ๐ ๐๐๐๐๐ข๐๐ก
๐ถ๐ = 3.5(0.01) + 0.84 = 0.875 ๐๐ฝ/๐๐ ๐
๐น๐ =132.04(0.05 โ 0.01)
0.45 โ 0.01= 12.00๐๐/โ
Thermal Requirements
๐๐ค๐ = 132.04(0.05 โ 0.0)[2500 โ (4.2 โ 1.90)(65)]
๐๐ค๐ โ12414.4๐๐ฝ
โร
1โ
3600๐ โ 3.45๐๐
๐๐ โ = 132.04(0.875 + 0.05(4.2))(65 โ 25)
๐๐ โ =5730.547
โร
1โ
3600๐ โ 1.59๐๐
๐๐โ = 12[1.0 + (0.01)(1.90)](65 โ 25)
173
๐๐โ โ407.6
โร
1โ
3600๐ โ 0.135 ๐๐
Total Thermal Requirement for drying
๐ = 3.45 + 1.59 + 0.135 = 5.18๐๐
174
APPENDIX B
PIPING AND FRICTION LOSSES CALCULATIONS
PIPE FROM RASPER TO MIXING TANK
Diameter of pipe, d = 82.9 mm
Length of pipe, L = 2m
Density of starch pulp, = 893 Kg/๐3
Velocity of the laminar flow ๐ = 2.0 m/s
Viscosity of starch at 25 โ, ๐ = 3.5 ร 10โ3 Pas
Specific gravity, ๐๐=
PRESSURE DROP
โ๐ =๐๐๐2๐๐
2๐
๐ต๐ข๐ก ๐ =64
๐ ๐
๐ต๐ข๐ก ๐ ๐ =๐ฃ๐
=
893 ร 2 ร 82.9 ร 10โ3
3.5 ร 10โ3= 42302.7
๐ =64
42302.7= 1.5 ร 10โ3
175
๐โ๐๐๐๐๐๐๐ โ๐ =1.5 ร 10โ3 ร 2 ร 4 ร 893 ร 0.893
2 ร 82.9 ร 10โ3= 57.72 ๐๐
MIXING TANK TO EXTRACTOR
Diameter of pipe, d = 82.9 mm
Length of pipe, L = 5m
Density of starch pulp, = 826 Kg/๐3
Velocity of the laminar flow ๐ = 2.0 m/s
Viscosity of starch at 25 โ, ๐ = 3.5 ร 10โ3 Pa.s
Specific gravity, ๐๐= 0.826
PRESSURE DROP
โ๐ =๐๐๐2๐๐
2๐
๐ต๐ข๐ก ๐ =64
๐ ๐
๐ต๐ข๐ก ๐ ๐ =๐๐
=
826 ร 2 ร 82.9 ร 10โ3
3.5 ร 10โ3= 39128.8
176
๐ =64
39128.8= 1.64 ร 10โ3
๐โ๐๐๐๐๐๐๐ โ๐ =1.64 ร 10โ3 ร 5 ร 4 ร 826 ร 0.826
2 ร 82.9 ร 10โ3= 134.97 ๐๐
FROM GROUP OF EXTRACTORS TO GROUP OF HYDROCYCLONES AND TO
GROUP OF SEPARATORS
Diameter of pipes, d = 66.7 mm
Total length of pipes, L = 4 m
Density of starch pulp, = 826 Kg/๐3
Velocity of the laminar flow ๐ = 2.0 m/s
Viscosity of starch at 25 โ, ๐ = 3.5 ร 10โ3 Pa.s
Specific gravity, ๐๐= 0.826
PRESSURE DROP
โ๐ =๐๐๐2๐๐
2๐
๐ต๐ข๐ก ๐ =64
๐ ๐
177
๐ต๐ข๐ก ๐ ๐ =๐๐
=
826 ร 2 ร 66.7 ร 10โ3
3.5 ร 10โ3= 62941.2
๐ =64
62941.2= 1.02 ร 10โ3
๐โ๐๐๐๐๐๐๐ โ๐ =1.02 ร 10โ3 ร 4 ร 4 ร 826 ร 0.826
2 ร 66.7 ร 10โ3= 83.47 ๐๐
FROM SEPARATOR TO DEWATERING EQUIPMENT
Diameter of pipe, d = 82.9 mm
Length of pipe, L = 6m
Density of starch pulp, = 893 Kg/๐3
Velocity of the laminar flow ๐ = 2.0 m/s
Viscosity of starch at 25 โ, ๐ = 3.5 ร 10โ3 Pa.s
Specific gravity, ๐๐= 0.893
PRESSURE DROP
โ๐ =๐๐๐2๐๐
2๐
178
๐ต๐ข๐ก ๐ =64
๐ ๐
๐ต๐ข๐ก ๐ ๐ =๐๐
=
893 ร 2 ร 82.9 ร 10โ3
3.5 ร 10โ3= 42302.7
๐ =64
42302.7= 1.51 ร 10โ3
๐โ๐๐๐๐๐๐๐ โ๐ =1.51 ร 10โ3 ร 6 ร 4 ร 893 ร 0.893
2 ร 82.9 ร 10โ3= 174.30 ๐๐
179
PIPING AND FRICTION LOSSES CALCULATIONS (GLUCOSE PLANT)
PIPE FROM MIXING TANK TO LIQUEFATCION VESSEL
Diameter of pipe, d = 84.7 mm
Length of pipe, L = 8m
Density of starch pulp, = 810 Kg/๐3
Velocity of the laminar flow ๐ = 2.0 m/s
Viscosity of starch at 25 โ, ๐ = 3.5 ร 10โ3 Pas
Specific gravity, ๐๐= 0.81
PRESSURE DROP
โ๐ =๐๐๐2๐๐
2๐
๐ต๐ข๐ก ๐ =64
๐ ๐
๐ต๐ข๐ก ๐ ๐ =๐ฃ๐
=
810 ร 2 ร 84.7 ร 10โ3
3.5 ร 10โ3= 39204
๐ =64
39204= 1.6 ร 10โ3
๐โ๐๐๐๐๐๐๐ โ๐ =1.6 ร 10โ3 ร 8 ร 4 ร 810 ร 0.81
2 ร 84.7 ร 10โ3= 198.30 ๐๐
180
PIPE FROM LIQUEFACTION VESSEL TO SACCHARIFICATION VESSEL
Diameter of pipe, d = 45 mm
Length of pipe, L = 12m
Density of starch pulp, = 1540 Kg/๐3
Velocity of the laminar flow ๐ = 2.0 m/s
Viscosity of glucose at 25 โ, ๐ = 6.8 ร 10โ3 Pa.s
Specific gravity, ๐๐= 1.54
PRESSURE DROP
โ๐ =๐๐๐2๐๐
2๐
๐ต๐ข๐ก ๐ =64
๐ ๐
๐ต๐ข๐ก ๐ ๐ =๐ฃ๐
=
1540 ร 2 ร 45 ร 10โ3
6.8 ร 10โ3= 20382
๐ =64
20382= 3.1 ร 10โ3
๐โ๐๐๐๐๐๐๐ โ๐ =3.1 ร 10โ3 ร 12 ร 4 ร 1540 ร 1.54
2 ร 45 ร 10โ3= 3921 ๐๐
181
PIPE FROM SACCHARIFICATION VESSEL ACIDIFICATION VESSEL
Diameter of pipe, d = 54.7 mm
Length of pipe, L = 10m
Density of starch pulp, = 1540 Kg/๐3
Velocity of the laminar flow ๐ = 2.0 m/s
Viscosity of glucose at 25 โ, ๐ = 6.8 ร 10โ3 Pas
Specific gravity, ๐๐= 1.54
PRESSURE DROP
โ๐ =๐๐๐2๐๐
2๐
๐ต๐ข๐ก ๐ =64
๐ ๐
๐ต๐ข๐ก ๐ ๐ =๐ฃ๐
=
1540 ร 2 ร 54.7 ร 10โ3
6.8 ร 10โ3= 24776
๐ =64
24776= 2.6 ร 10โ3
๐โ๐๐๐๐๐๐๐ โ๐ =2.6 ร 10โ3 ร 10 ร 4 ร 1540 ร 1.54
2 ร 54.7 ร 10โ3= 36272 ๐๐
182
PIPE FROM ACIDIFICATION TANK TO ADSORPTION COLUMN
Diameter of pipe, d = 30.1 mm
Length of pipe, L = 15m
Density of starch pulp, = 1540 Kg/๐3
Velocity of the laminar flow ๐ = 2.0 m/s
Viscosity of glucose at 25 โ, ๐ = 6.8 ร 10โ3 Pas
Specific gravity, ๐๐= 1.54
PRESSURE DROP
โ๐ =๐๐๐2๐๐
2๐
๐ต๐ข๐ก ๐ =64
๐ ๐
๐ต๐ข๐ก ๐ ๐ =๐ฃ๐
=
1540 ร 2 ร 30.1 ร 10โ3
6.8 ร 10โ3= 13634
๐ =64
13634= 4.7 ร 10โ3
๐โ๐๐๐๐๐๐๐ โ๐ =4.7 ร 10โ3 ร 15 ร 4 ร 1540 ร 1.54
2 ร 30.1 ร 10โ3= 11096 ๐๐
183
PIPE FROM ADSORPTION COLUMN TO ION EXCHANGE COLUMN (CATION)
Diameter of pipe, d = 54.7 mm
Length of pipe, L = 15m
Density of starch pulp, = 1540 Kg/๐3
Velocity of the laminar flow ๐ = 2.0 m/s
Viscosity of glucose at 25 โ, ๐ = 6.8 ร 10โ3 Pas
Specific gravity, ๐๐= 1.54
PRESSURE DROP
โ๐ =๐๐๐2๐๐
2๐
๐ต๐ข๐ก ๐ =64
๐ ๐
๐ต๐ข๐ก ๐ ๐ =๐ฃ๐
=
1540 ร 2 ร 54.7 ร 10โ3
6.8 ร 10โ3= 24776
๐ =64
24776= 2.6 ร 10โ3
๐โ๐๐๐๐๐๐๐ โ๐ =2.6 ร 10โ3 ร 15 ร 4 ร 1540 ร 1.54
2 ร 54.7 ร 10โ3= 3381 ๐๐
184
PIPE FROM TO ION EXCHANGE COLUMN (CATION) TO ION EXCHANGE
COLUMN (ANION)
Diameter of pipe, d = 45 mm
Length of pipe, L = 6 m
Density of starch pulp, = 1540 Kg/๐3
Velocity of the laminar flow ๐ = 2.0 m/s
Viscosity of glucose at 25 โ, ๐ = 6.8 ร 10โ3 Pa.s
Specific gravity, ๐๐= 1.54
PRESSURE DROP
โ๐ =๐๐๐2๐๐
2๐
๐ต๐ข๐ก ๐ =64
๐ ๐
๐ต๐ข๐ก ๐ ๐ =๐ฃ๐
=
1540 ร 2 ร 45 ร 10โ3
6.8 ร 10โ3= 20382
๐ =64
20382= 3.1 ร 10โ3
๐โ๐๐๐๐๐๐๐ โ๐ =3.1 ร 10โ3 ร 8 ร 4 ร 1540 ร 1.54
2 ร 45 ร 10โ3= 2614 ๐๐
185
LENGTH OF PIPING IN THE SORBITOL PLANT
PIPE FROM DEXTROSE STORAGE TANK TO HYDROGENATION VESSEL
DATA
Internal diameter of pipe, d = 108.3mm
Density of glucose, = 1540 kg/m3
Velocity of the laminal flow = 2.0 m/s
Pipe length, l = 3.0m
Viscosity of glucose = 0.0068 Pa.s
Specific gravity, Sg = 1.54
PRESSURE DROP
โ๐ =๐๐๐2๐๐
2๐
๐ต๐ข๐ก ๐ =64
๐ ๐
๐ต๐ข๐ก ๐ ๐ =๐๐
= 9060
๐ =64
9060= 7.06 ร 10โ3
186
๐โ๐๐๐๐๐๐๐ โ๐ =7.06 ร 10โ3 ร 3 ร 4 ร 1540 ร 1.54
2 ร 108.3 ร 10โ3= 0.93๐๐
PIPE FROM HYDROGENATION REACTOR TO EVAPORATOR
DATA
Internal diameter of pipe, d = 135.7mm
Density of glucose, = 1490 kg/m3
Velocity of the laminal flow = 2.0 m/s
Pipe length, l = 3m
Viscosity of sorbitol, = 0.110 Pa.s
Specific gravity, Sg = 1.49
187
PRESSURE DROP
โ๐ =๐๐๐2๐๐
2๐
๐ต๐ข๐ก ๐ =64
๐ ๐
๐ต๐ข๐ก ๐ ๐ =๐ฃ๐
= 867
๐ =64
867= 0.0738
๐โ๐๐๐๐๐๐๐ โ๐ =0.0738 ร 3 ร 4 ร 1490 ร 1.49
2 ร 135.7 ร 10โ3= 7.2 ๐๐๐
EVAPORATOR TO CRYSTALLISER
DATA
Internal diameter of pipe, d = 134.5mm
Density of glucose, = 1490 kg/m3
Velocity of the laminal flow = 2.0 m/s
Pipe length, l = 3.0 m
Viscosity of sorbitol, = 0.110 Pa.s
Specific gravity, Sg = 1.49
188
PRESSURE DROP
โ๐ =๐๐๐2๐๐
2๐
๐ต๐ข๐ก ๐ =64
๐ ๐
๐ต๐ข๐ก ๐ ๐ =๐ฃ๐
= 867
๐ =64
867= 0.0738
๐โ๐๐๐๐๐๐๐ โ๐ =0.0738 ร 3 ร 4 ร 1490 ร 1.49
2 ร 134.5 ร 10โ3= 7.3 ๐๐
189
APPENDIX C
DETAILED EQUIPMENT DESIGN CALCULATION
TOP SUSPENDED MOTOR CYLINDRICAL-SCREEN CENTRIFUGE
CALCULATIONS
CENTRIFUGAL FORCE
๐ถ๐๐๐ก๐๐๐๐ข๐๐๐ ๐๐๐๐๐, ๐น๐ = ๐๐๐2
Where
๐น๐= centrifugal force generated
๐ = mass of rotating particle (Kg)
๐ = angular speed (rad/min)
๐ = radius of rotation (m)
๐น๐ = ๐๐๐2 =๐๐ข๐ก
2
๐
๐ข๐ก =2๐๐๐
60
Hence
๐น๐ = ๐
๐(2๐๐๐
60)2
190
But
๐ = total mass of basket content (magma) = 1000 kg
๐ = speed of rotation (rev/min) = 1200 rev/min
๐ = radius of rotation (m) = 0.6m
๐น๐ = ๐
๐(2๐๐๐
60)2
= 1000
0.6(2๐ ร 0.6 ร 1200
60)
2
= 9474.8 ๐พ๐
SEPARATION FACTOR
The separation factor is an important characteristic of centrifuges, which is used to determine the
settling rate of particles in a field of centrifugal force. This is achieved from classical equations
by replacing the Archimedes number Ar by (Ar)(Kp). The separation factor can also be used as a
basis for classifying centrifuges into normal which has a Kp < 3500 and ultracentrifuges with the
Kp > 3500.
๐ฒ๐ท =๐๐
๐๐
๐พ๐ = separation factor
๐ = angular speed (rad/sec)
๐ = acceleration due to gravity (m/๐ 2) = 9.81 m/๐ 2
๐ = radius of rotation = 0.6m
๐ฒ๐ท =๐๐
๐๐
191
But ๐ =2๐๐
60=
2๐ร1200
60= 40๐
๐ฒ๐ท =(40๐)2
9.81ร0.6= 2682.9 โ 2683
CAPACITY FACTOR
Another important characteristic of centrifuges is the capacity factor CF, defined as the product
of the area of the cylindrical surface available for collection of sediments A, and the separation
factor ๐ฒ๐ท.
๐ช๐ญ = ๐จ ร ๐ฒ๐ท
But ๐จ =๐ (๐ซโ๐)
๐ฏ
Where
๐ซ = diameter of basket = 1.2m
๐ = thickness of fluid on basket surface = 0.14m
๐ฏ = length of cylindrical surface in contact with fluid = 1.0m
(๐ซ โ ๐) = average diameter of rotation =๐ซ+(๐ซโ๐๐)
๐
Hence
๐จ =๐ (๐ซโ๐)
๐ฏ=
๐ ร(๐.๐โ๐.๐๐)
๐= ๐. ๐๐๐๐
192
๐ช๐ญ = ๐. ๐๐ ร 2683 = 8934.39
= 8934.39 ๐2
CAPACITY OF CENTRIFUGE
๐๐ก = ๐๐ ร ๐ถ๐น
๐๐ =1
18
๐๐๐(๐๐โ๐๐)
๐
๐ ๐ = Reynolds number
๐ = viscosity of fluid = 199 Pa.s
๐๐ = diameter of particle = 0.002mm
๐๐ =149 Kg/๐3
๐ = 9.81 m/๐ 2
๐๐ = density of fluid = 1586.2 Kg/๐3
๐๐ =1
18
0.002 ร 103 ร 9.81 ร (1586.2 โ 149)
199= 7.9๐/๐
๐๐ก = ๐๐ ร ๐ถ๐น
= 7.9 ร 8934.39 = 70581.7 ๐3/๐
Capacity of centrifuge is 70581.7 ๐3/๐
193
STRESS IN THE INTERNAL BASKET
The centrifugal force generated exerts stress on the internal basket. Hence it is very vital to
consider the strength of the material used for fabrication of the internal basket of the centrifuge.
They are expected to be adequately strong enough to withstand the stress during operation in
order to avoid possible basket failure. The formula for determining the stress per unit area on the
wall of the basket established by Kreg (1975) was utilized;
๐๐ =๐๐๐2๐
๐๐ท๐ป=
๐๐๐2
๐๐ป
Where
๐๐ = stress on the walls of internal basket (N/๐2)
๐๐ = total mass of basket and its content (Kg) = 2110 Kg
๐ป = height of basket (m) = 1.0m
๐ท = Diameter of basket (m) = 1.2 m
๐ = 40๐
๐๐ =2110 ร (40๐)2
๐ ร 1= 10606 ๐พ๐
๐๐ = 10606 ๐พ๐
194
ALLOWABLE THICKNESS OF MATERIAL FOR THE CONSTRUCTION OF THE
INTERNAL BASKET
The allowable thickness of the material for construction internal basket is the minimum
thickness of the basket material that can withstand the expected stresses to be exerted on the
walls of the basket in other to prevent avoidable basket failures. Kreg (1976) stated that the
thickness of the wall of the basket to withstand the stress is a function of the unit stress that acts
on the wall, the diameter of the basket and the maximum permissible stress of the material as
shown below:
๐ก๐ =๐๐๐ท
2๐๐
Where
๐ก๐= allowable thickness of basket material
๐๐ =10606 ร 103
๐๐= permissible stress of the material of the basket = 115 ร 106 ๐/๐2
๐ท = Diameter of basket (m) = 1.2m
๐ก๐ =๐๐๐ท
2๐๐=
10606 ร 103 ร 1.2
2 ร 115 ร 106= 0.06๐
๐ก๐ = 0.06๐
195
POWER REQUIREMENT
The power required to drive the machine (internal basket) is a function of the mass of the internal
basket, its content, flanges and the central shaft that transmits power from the electric motor to
the basket through pulleys and belts as shown in Figs 5 and 7. Hence the power required to rotate
or drive the basket for separation of the sugar crystals is obtained by using the generally
established equation.
๐๐๐ค๐๐ ๐๐๐๐ข๐๐๐๐๐๐๐ก = ๐น๐ ร ๐
Where
๐น๐ = total force of basket, shaft and basket content (N)
๐น๐ = ๐๐๐2๐
Assuming ๐๐ = 1000 + 1110 = 2110 ๐พ๐
๐น๐ = 2110 ร (40๐)2 ร 0.6 = 19991.9 ๐พ๐
๐ = velocity of basket at full speed (m/sec) = ๐๐
๐ = 40๐ ร 0.6 = 24๐ = 75.4๐
๐ ๐๐
๐๐๐ค๐๐ ๐๐๐๐ข๐๐๐๐๐๐๐ก = 75.4 ร 19991.9 = 1507389.26 ๐พ๐
196
TWISTING MOMENT
The high rotating speed of the shaft which is attached to the internal basket is subjected to a
twisting moment. In order for the shaft not to fail, the value of the twisting moment generated is
expected to be within the permissible limit in order to avoid failure of the shaft. The expression
of Juvinall (1976) was used to determine the expected twisting moment of the shaft as shown
below.
Where
๐๐ก = twisting movement (Nm)
W = power transmitted (watts)
N = speed of rotation of the shaft (rev/sec)
๐๐ก =60 ร 1507389.26 ร 103
2๐ ร120060
= 719725.4 ๐พ๐๐
DIAMETER OF SHAFT
The minimum diameter of the shaft to transmit power to the internal basket is dependent on the
twisting moment (Torque) on the shaft and the permissible shear stress of the material used to
make (stainless steel) the shaft as shown (Holman 1969).
๐ =(16๐๐ก๐ท)0.33
๐๐๐
197
๐ = diameter of shaft
๐๐ก = twisting moment on shaft due to rotation of the internal basket fastened on shaft.
๐ท = diameter of basket
๐๐ = permissible shear stress of stainless steel = 115 ร 106 ๐/๐2
๐ =(16ร719725.4ร103ร1.2)0.33
๐ร115ร106= 6.1 ร 10โ6m
198
ADSORPTION COLUMN DESIGN CALCULATIONS
FIXED BED DESIGN PARAMETERS
Glucose syrup flowing at a rate, G02 = 5104 kg/day through the column.
Density of glucose =1540 kg/m3
๐บ02 = 5104๐๐
๐๐๐ฆ
The volumetric flow rate ๐ =๐๐๐ ๐ ๐๐๐๐ค ๐๐๐ก๐
๐๐๐๐ ๐๐ก๐ฆ
=5104
1540
= 3.3 ๐3
๐๐๐ฆ
= 3314 ๐ฟ/๐๐๐ฆ
The concentration of the colorants, ashes, and other impurities in the 3314 ๐ฟ/๐๐๐ฆ is 7.6 ๐๐/๐ฟ .
Assuming 90% removal, then the concentration remained in the flow stream is 0.76 ๐ ๐/๐ฟ
Liquefaction time is 90 minutes, and then the following parameters are chosen from above table
The Lagmuir isotherm parameter for NORIT; ๐๐, K and R2 values
Liquefaction
Time
๐๐ K R2
45 80.65 18.87 0.94
60 179.57 37.99 0.90
75 277.78 46.69 0.96
90 344.83 26.93 0.94
NORIT is a type of Activated Carbon to be used in this project
199
๐๐ = 344.83 ๐๐/๐
K=26.93 L/mg
๐ถ๐ = 0.76๐๐
๐ฟ
๐ถ๐
๐๐=
1
๐๐๐พ1+
๐ถ๐
๐๐
๐๐ = ๐๐ ๐กโ๐ ๐๐๐ ๐ ๐๐ ๐๐๐ ๐๐๐๐๐ก๐ ๐๐๐ ๐๐๐๐๐ ๐๐๐ ๐ข๐๐๐ก ๐๐๐ ๐ ๐๐ ๐๐๐ ๐๐๐๐๐๐ก
๐ถ๐ = concentration of adsorbate remaining in the syrup after adsorption
๐พ1, ๐๐ are Langmuir constant representing the maximum adsorption capacity and energy of
adsorption respectively.
0.76
๐๐=
1
344.83 ร 26.93+
0.76
344.83
0.76
๐๐= 2.3 ร 10โ3
๐๐ = 328.8 ๐g/g Carbon (C)
Therefore 328.8 mg of adsorbate is adsorbed per every gram of adsorbent
Impurity Load = (๐ผ๐๐๐ก๐๐๐ ๐๐๐๐๐๐๐ก๐๐๐๐๐ โ ๐๐๐๐๐ ๐๐๐๐๐๐๐ก๐๐๐ก๐๐๐) ๐๐/๐ ร 3314 ๐ฟ/๐๐๐ฆ
Impurity Load = (7.6 โ 0.76) ๐๐/๐ ร 3314 ๐ฟ/๐๐๐ฆ
= 22667 mg/day
200
BED PARAMETERS
Diameter of adsorbent Bed= ๐ท๐ต๐๐
๐ด๐ต๐๐ =๐๐ท2
๐ต๐๐
4
=๐ ร1.52
4
= 1.8 ๐2
Calculating the volume of the adsorbent Bed
๐๐ต๐๐ = ๐ดBedโ HBed
For a bed height ๐ป๐ต๐๐ = 2 ๐
Volume of bed, VBed =๐ดBedโ HBed
VBed = 1.8 โ 2
VBed = 3.6 ๐3
The volume of the adsorbent bed to be used is 3.6 ๐3
201
COLUMN DESIGN SPECIFICATIONS AND PARAMETERS
The volume of the column is the sum of the volume of the adsorbent bed (๐๐ต๐๐) and the volume
of the liquid-filled void(๐๐ฟ).
๐๐ด๐๐ = ๐๐ต๐๐ + ๐๐ฟ
The adsorbent should cover at least 66% of the volume of the adsorbent column
๐๐ต๐๐ = 66% โ ๐๐ด๐๐
VBed = 3.6 ๐3
๐๐ด๐๐ = ๐๐ต๐๐
0.66=
3.6
0.66
๐๐ด๐๐ = 5.5 ๐3
๐๐ฟ = ๐๐ด๐๐ โ ๐๐ต๐๐
๐๐ฟ = 5.5 โ 3.6
๐๐ฟ = 1.9 ๐3
The volume of the liquid-filled void, ๐๐ฟis the volume occupied by the glucose syrup and is
between the bed surface and the rim of the column.
The diameter of the adsorbent bed is the same as the diameter of the column
๐ท๐ต๐๐ = ๐ท๐ด๐๐
๐๐ด๐๐ = ๐ด๐ด๐๐ โ ๐ป๐ด๐๐
๐ป๐ด๐๐ =4โ ๐๐ด๐๐
๐๐ท2๐ต๐๐
202
๐ป๐ด๐๐ =5.5โ4
๐โ 1.52
= 3 ๐
BED POROSITY, ๐บ
The bed porosity is the void fraction of the reactor volume. It is given by the equation below
๐ =๐๐๐๐ ๐ฃ๐๐๐ข๐๐(๐๐ฟ)
๐ด๐๐ ๐๐๐๐๐ ๐ฃ๐๐๐ข๐๐(๐๐ด๐๐ )=
๐๐ด๐๐ โ๐๐ต๐๐
๐๐ด๐๐ = 1 โ
๐๐ต๐๐
๐๐ด๐๐
= 1 โ3.6
5.5
= 0.3
๐โ๐ ๐๐๐๐๐ ๐๐ก๐ฆ ๐๐ ๐กโ๐ ๐๐๐ , ๐ ๐๐ 0.3
FILTRATION RATE
The above volumetric flow of the glucose 3.98 ๐3
๐๐๐ฆ syrup will be pumped for 30 minutes that is
1809 seconds
๐๐ด๐๐ =3.98 ๐3
1809 ๐
= 0.0022 ๐3
Therefore the filtration rate (FR) = ๐๐ด๐๐
๐ด๐ด๐๐
= 0.0022
1.8
= 0.001๐3
๐2๐
203
RESIDENCE TIME
According to the different definitions for the flow velocity, two different residence times can be
defined (Eckhard 2012). These are the Empty Bed Contact Time (EBCT), and the effective
contact time, ๐.
Empty Bed Contact Time (EBCT)
This is the residence time for the empty adsorption column.
๐ธ๐ต๐ถ๐ =๐๐ด๐๐
๐
=5.5
0.0022
= 2500 ๐
= 42 ๐๐๐๐ข๐ก๐๐
Effective Contact Time, ๐
This is defined as the quotient of the free bed volume available for liquid flow divided by the
crude glucose flow rate through the bed.
๐ =๐โ๐๐ด๐๐
๐ = ๐ธ๐ต๐ถ๐ ร ๐
= 42 ร 0.3
= 12.5 ๐๐๐
Therefore, the effective contact time, ๐ for a flow rate of 0.0022 ๐3
๐ ๐๐ 12.5 ๐๐๐utes
204
COSTING OF ADSORPTION COLUMN
The capital cost of the adsorption column takes into consideration
I. The cost of the column, C๐ด๐๐ given in Euros
II. The cost of adsorbent bed (fixed), C๐ต๐๐
THE COST OF THE COLUMN (๐๐จ๐ ๐)
The cost of the column is calculated using the following empirical equation
C๐ด๐๐ = 583.6 โ ๐ท0.675 โ ๐ป โ ๐น๐๐๐ก โ (๐โ145
50)0.44
๐ท = ๐กโ๐ ๐๐๐๐๐๐ก๐๐ ๐๐ ๐๐๐๐ข๐๐
๐ป = ๐กโ๐ โ๐๐๐โ๐ก ๐๐ ๐๐๐๐ข๐๐
๐ = ๐กโ๐ ๐ค๐๐๐๐๐๐ ๐๐๐๐ ๐ ๐ข๐๐
๐น๐๐๐ก = ๐๐๐๐๐๐ ๐๐๐ก๐ ๐ ๐๐๐๐๐๐๐ก๐๐๐ ๐ก๐ ๐ก๐๐๐ ๐๐๐ก๐ ๐๐๐๐๐ข๐๐ก ๐๐๐ ๐กโ๐ ๐๐๐ ๐ก ๐๐ ๐กโ๐ ๐๐๐ก๐๐๐๐๐
The parameters of the proposed design are
๐ท๐ด๐๐ = 1.5 ๐
๐ป๐ด๐๐ = 3 ๐
๐ = 100 ๐พ๐๐
๐น๐๐๐ก ๐๐๐ ๐ ๐ก๐๐๐๐๐๐ ๐ ๐ ๐ก๐๐๐ 304 ๐๐ 1.7
C๐ด๐๐ = 583.6 โ 1.50.675 โ 3 โ 1.7 โ (120โ145
50)0.44
= โฌ 51,385.6
C๐ด๐๐ = $ 57,551
205
THE COST OF THE FIXED BED(๐๐ฉ๐๐ )
The cost of the packed or fixed bed is evaluated from the following equation
C๐ต๐๐ =๐โ๐ท2
4โ ๐ป โ ๐ถ
Where
๐ป = ๐กโ๐ โ๐๐๐โ๐ก ๐๐ ๐ต๐๐
D = ๐กโ๐ โ๐๐๐โ๐ก ๐๐ ๐ต๐๐
C= the volume cost of the packing (Bed). The cost of 1m3 of activated carbon is
3500 EUR which is about 3908.45 USD
๐ท๐ต๐๐=
๐ป๐ต๐๐
C๐ต๐๐ =๐โ๐ท2
4โ ๐ป โ ๐ถ
=๐โ1.52
4โ 2 โ 3500
= โฌ12,372
C๐ต๐๐ = $13856
The total cost of the Adsorption column unit, CT = C๐ด๐๐ + C๐ต๐๐
=(57,551 + 13856 ) $
CT= $71,407
Therefore the total cost of the Adsorption column is $71,407
206
HYDROGENATION REACTOR DESIGN
DESIGN EQUATIONS AND DIMENSIONS OF HYDROGENATION REACTOR
REACTOR DESIGN
VOLUME OF THE TANK
Since the shape of the reactor is cylindrical, the volume (V) is calculated by the formula below.
๐ = ๐๐ 2๐ป
๐โ๐๐๐ ๐ = ๐๐๐๐ข๐๐ ๐๐ ๐กโ๐ ๐๐๐ก๐๐๐ ๐๐๐๐๐ก๐๐
๐ = ๐ ๐๐๐๐ข๐ ๐๐ ๐กโ๐ ๐๐๐๐๐ก๐๐ = 0.9๐
๐ป = ๐ป๐๐๐โ๐ก ๐๐ ๐กโ๐ ๐๐๐๐๐ก๐๐ = 4.0๐
๐ท๐๐๐๐๐ก๐๐ ๐๐ ๐กโ๐ ๐๐๐๐๐ก๐๐, ๐ท = 1.8๐
๐โ๐๐๐๐๐๐๐ ๐ ๐๐๐๐ข๐ , ๐ = ๐ท
2
๐ ๐๐๐๐ข๐ = 1.8
2 = 0.9๐
๐ป๐๐๐๐ ๐ฃ๐๐๐ข๐๐ ๐๐ ๐กโ๐ ๐ก๐๐๐, ๐ = ๐ ร (0.9)2 ร 4.0
๐ = 10.2๐3
207
VOLUME OF REACTING CHAMBER (๐ฝ๐น)
Still using the equation for the volume of a cylinder,
๐๐ = ๐ ๐2โ
๐โ๐๐๐ ๐ = ๐ ๐๐๐๐ข๐ ๐๐ ๐กโ๐ ๐๐๐๐๐ก๐๐๐ ๐โ๐๐๐๐๐ = 0.80๐
โ = โ๐๐๐โ๐ก ๐๐ ๐กโ๐ ๐๐๐๐๐ก๐๐๐ ๐๐โ๐๐๐๐ = 3.8๐
๐ท๐๐๐๐๐ก๐๐ ๐๐ ๐กโ๐ ๐๐๐๐๐ก๐๐๐ ๐โ๐๐๐๐๐ (๐) = 1.6๐
๐ ๐๐๐๐ข๐ ๐ = ๐
2
๐ ๐๐๐๐ข๐ ๐ = 1.6
2 = 0.80๐
๐๐๐๐ข๐๐ ๐๐ ๐กโ๐ ๐๐๐๐๐ก๐๐๐ ๐โ๐๐๐๐๐ , ๐๐ = ๐ ร (0.802) ร 1.8
๐๐ โ 7.64๐3
STIRRER DESIGN
According to Davis (2009), for a cylindrical stirred tank, the ratio of the diameter of the
impeller(DI) to the diameter of the tank (DT) is within the range
0.3 <๐ท๐ผ
๐ท๐< 12
Using similar conditions for the calculation of the diameter of the impellers, Assuming
๐ท๐ผ
๐ท๐ = 0.3
๐โ๐๐๐ ๐ท๐ผ = ๐ท๐๐๐๐๐ก๐๐ ๐๐ ๐กโ๐ ๐๐๐๐๐๐๐๐
208
๐ท๐ = ๐ท๐๐๐๐๐ก๐๐ ๐๐ ๐กโ๐ ๐ก๐๐๐ = 1.8๐
๐ท๐ผ = 0.3๐ท๐ = 0.3 ร 1.8
๐ท๐ผ = 0.54 ๐.
NUMBER OF IMPELLERS (N)
According to Davis (2009), the number of impellers N is given by
๐ป โ ๐ท๐ผ
๐ท๐ผ > ๐ >
๐ป โ 2๐ท๐ผ
2๐ท๐ผ
๐โ๐๐๐ ๐ป = ๐กโ๐ โ๐๐๐โ๐ก ๐๐ ๐กโ๐ ๐ก๐๐๐ = 3.0 ๐
๐๐ข๐ก ๐ป โ 2๐ท๐ผ
2๐ท๐ผ =
3 โ 2(0.42)
2(0.42)
๐ โ 3
๐โ๐๐๐๐๐๐๐ ๐กโ๐ ๐๐ข๐๐๐๐ ๐๐ ๐๐๐๐๐๐๐๐๐ ๐๐ 3
IMPELLER TIP SPEED (n)
๐ = ๐ ร ๐ ร ๐ท๐
๐ = ๐ ร 3 ร 1.4
๐ = 13.19๐๐ โ1
๐โ๐๐ ๐๐ ๐๐๐ข๐๐ฃ๐๐๐๐๐ก ๐ก๐ 180 ๐๐๐
209
FLOW OF FLUID IN THE REACTOR
To identify the flow of the fluid in the reactor, the Reynolds number needs to be calculated.
๐๐ = ๐2๐๐
๐
๐คโ๐๐๐ ๐ = ๐กโ๐ ๐๐๐๐๐๐ก๐๐ ๐๐ ๐กโ๐ โ๐ฆ๐๐๐๐๐๐๐๐ก๐๐๐ ๐๐๐๐๐ก๐๐ = 1.8๐
๐ = ๐กโ๐ ๐๐๐๐๐๐๐๐ ๐ก๐๐ ๐ ๐๐๐๐ = 13.19๐๐ โ1
๐ = ๐๐๐๐ ๐๐ก๐ฆ ๐๐ ๐ ๐๐๐๐๐ก๐๐ = 1490 ๐๐๐โ3
๐ = ๐กโ๐ ๐ฃ๐๐ ๐๐๐ ๐๐ก๐ฆ ๐๐ ๐ ๐๐๐๐๐ก๐๐ ๐๐๐๐๐ข๐๐๐ = 0.11 ๐๐๐
๐๐ = (1.8)2 ร 13.19 ร 1490
0.11
๐๐ = 578873.12
๐โ๐๐ ๐๐ ๐๐๐๐๐ก๐๐ ๐กโ๐๐ 10000 (๐๐
> 10000)๐๐๐ ๐กโ๐๐๐๐๐๐๐ ๐กโ๐ ๐๐๐๐ค ๐๐๐ ๐๐๐ ๐กโ๐ ๐ก๐๐๐ ๐๐ ๐ก๐ข๐๐๐ข๐๐๐๐ก.
ELECTRONIC MOTOR REQUIREMENT
A 5-pole 3HP AC motor operating at a frequency of 10 Hz with a controllable rotor speed of
200rpm will be used.
๐ = ๐๐ข๐๐๐๐ ๐๐ ๐๐๐๐๐ = 2
๐ = ๐๐๐ก๐๐ ๐๐๐๐๐๐ก๐๐๐ ๐๐๐๐๐ข๐๐๐๐ฆ
๐ = ๐ ๐๐๐๐ ๐๐ ๐๐๐ก๐๐ = 200 ๐๐๐
210
๐๐ = ๐๐๐ก๐๐ ๐๐๐ค๐๐ = 3 โ๐๐๐ ๐ ๐๐๐ค๐๐ (๐ป๐)
SYNCHRONOUS SPEED (NS)
This refers to the rotation speed of the stators magnetic field.
๐๐ = 120 ร ๐
๐
๐๐ = 120 ร 10
5
๐๐ = 240 ๐๐๐
TORQUE (T)
This is the force required to achieve one complete revolution
๐ = ๐๐ ร 5252
๐๐
๐ = 3 ร 5252
240
๐ = 65.65 ๐ผ๐๐๐ก
VOLUME OF CATALYST BED (V)
For the efficient conversion of glucose to sorbitol, the volumetric flow rate ๐๐ must be between
0.5- 3.5 volume of feed/ volume of catalyst.
But volume flow rate of feed of feed = ๐๐ = 0.2247 ๐3/๐๐๐
๐กโ๐ ๐ฃ๐๐ข๐๐ ๐๐ ๐๐๐๐ ๐๐๐ โ๐๐ข๐ = 13.482๐3/โ
๐๐
๐๐ถ= 0.5
๐๐๐ข๐๐ ๐๐ ๐๐๐ก๐๐๐ฆ๐ ๐ก = 0.5๐๐
211
๐๐๐ข๐๐ ๐๐ ๐๐๐ก๐๐๐ฆ๐ ๐ก = 0.5 ร 13.482 = 6.74 ๐3
AREA OF CATALYST ON BED (A)
The height of the catalyst bed = 0.2m
๐ด๐๐๐ ๐๐ ๐๐๐ก๐๐๐ฆ๐ ๐ก = 6.74
0.2= 33.7 ๐2