1

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 ©2016 University of Ghana

All rights reserved

vacharibasam@gmail.com

obirida@gmail.com

francisopoku00@gmail.com

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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: ……………………………………………..

Date: …………………………………………………..

STUDENT NAME: ANANEY-OBIRI DANIEL

Signature: …………………………………………….

Date: …………………………………………………..

STUDENT NAME: POKU FRANCIS

Signature: …………………………………………….

Date: …………………………………………………..

SUPERVISOR: DR. GEORGE AFRANE

Signature: ……………………………………………

Date: …………………………………………………..

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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 plant’s service life is estimated to be 20 years with an initial capital investment and working

capital of GHȻ 4,754,314 and GHȻ 838,997 respectively. The Net Present Value (NPV) of the

project from the sensitivity analysis is GHȻ2,783,058 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 GHȻ 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 (𝐂𝐌) ........................................................ 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 Cato’s 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 in‐terms 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 (𝐶6𝐻12𝑂6) 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 D‐sorbitol are synthesized each year worldwide by the

catalytic hydrogenation of D‐glucose. This represents an inexpensive, abundant feed‐ stock

obtainable from renewable resources such as starch‐ 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 (140–200 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 (60–70% moisture content and 45–55%, 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

12

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 °Bé. Water is then

separated from the starch slurry, increasing the concentration to 18 to 20 °Bé, 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-

170°C 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 – 95°C (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, α-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, α-amylase. Calcium

ions, about 50ppm quantities are added to stabilise the enzyme in the Continuous Stirring

Reaction Tank (CSRT). The enzyme hydrolyses the α-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 α-amylase and for optimal performance

of the enzyme glucoamylase. The glucoamylase is added to hydrolyse both the α-1,4 linkage and

the α-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

(𝐶6𝐻10𝑂5)𝑛 + 𝑛𝐻2𝑂 𝑔𝑙𝑢𝑐𝑜𝑎𝑚𝑦𝑙𝑎𝑠𝑒 𝑛𝐶6𝐻12𝑂6

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 150–170°F (69–77°C) with a typical contact

time of 90–120 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

∆𝐸 = ��(∆𝐻 + ∆𝐾𝐸 + ∆𝑃𝐸) ± 𝑄 ± 𝑊

At steady state, ∆𝐸 = 0

There is no movement of equipment, velocity is zero hence ∆𝐾𝐸 = 0

∆𝑃𝐸 = 0

The above equation reduces to

𝑄 = ��∆𝐻 + 𝑊

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 ×

FT)

TOTAL AREA

(FT2)

SECURITY CHECKPOINT (1,2 & 3) 5×5 EACH 75

RAW MATERIAL RECEPTION 10×15 150

CHANGING ROOM 25×30 750

WASHROOM 19 ×25 475

CANTEEN 19×32 608

EXPANSION AREA 100×200 20,000

VISITORS CAR PARK 18×25 450

WORKERS CAR PARK 19×25 475

LOADING TRUCKS PARK (1 & 2) 19×25 475

RECEPTION 8×10 80

CLINIC 19×22 418

WAREHOUSE 1 & 2 70×100 7,000

ADMINISTRATION BLOCK 27×38 1,026

QUALITY ASSURANCE DEPARTMENT 25×30 750

ENGINEERING DEPARTMENT 19×25 475

UTILITIES 19×25 475

STARCH EXTRACTION AREA 66×115 7,590

DEXTROSE PREPARATION PLANT 60×100 6,000

SORBITOL PRODUCTION AREA 60×100 6,000

WASTE TREATMENT PLANT 19×22 418

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 – 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

× 10−5

Mixing tank Extractor 10 5 82.9 1.35 × 10−4

Extractor Hydrocyclone

and separator

160 4 66.7 8.347

× 10−5

Separator Dewatering

machine

10 6 82.9 1.74 × 10−4

Mixing tank Liquefaction

vessel

5 8 84.7 1.98 × 10−4

Liquefaction

vessel

Saccharification

vessel

5 12 45 3.92 × 10−3

Saccharification

vessel

Acidification

tank

10 10 54.7 2.28 × 10−3

48

Acidification

tank

Adsorption

column

5 15 30.1 0.011

× 10−3

Adsorption

column

Cation exchange

column

10 15 54.7 3.39 × 10−3

Cation

exchange

Anion exchange

column

5 8 45 2.6× 10−3

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, 𝜔 is the

angular velocity and Fc is the centrifugal force generated.

𝐹𝑐 = 𝑚𝑟𝜔2 ………………………………… . .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 (𝒎𝟐) 8934.39

Theoretical capacity of centrifuge 𝑸𝒕 (𝒎𝟑/𝒔) 70581.7

Stress in the internal Basket 𝜹𝒃(𝒌𝑵) 10606

Allowable thickness of material for

construction of internal basket 𝒕𝒃(𝒎)

0.06

Power requirement P (kW) 1507389.2

Twisting moment 𝑴𝒕(𝒌𝑵) 719725

Minimum diameter of shaft (𝒎) 6. × 𝟏𝟎−𝟔

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 “adsorb” 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

“suspended” 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

𝐶𝑒

𝑞𝑒=

1

𝑞𝑜𝐾1+

𝐶𝑒

𝑞𝑜

𝑞𝑒 = 𝑖𝑠 𝑡ℎ𝑒 𝑚𝑎𝑠𝑠 𝑜𝑓 𝑎𝑑𝑠𝑜𝑟𝑏𝑎𝑡𝑒 𝑎𝑑𝑠𝑜𝑟𝑏𝑒𝑑 𝑝𝑒𝑟 𝑢𝑛𝑖𝑡 𝑚𝑎𝑠𝑠 𝑜𝑓 𝑎𝑑𝑠𝑜𝑟𝑏𝑒𝑛𝑡

𝐶𝑒 = concentration of adsorbate remaining in the syrup after adsorption

𝐾1, 𝑞𝑜 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.

• Column (Adsorber)

• Adsorbent Bed (GAC)

• Effluent Pipe

• Influent Pipe

• 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 𝑄𝑎𝑑 𝑚3/𝑑𝑎𝑦 3.3

Fixed-bed height 𝐻𝐵𝑒𝑑 𝑚 2

Fixed-bed diameter 𝐷𝐵𝑒𝑑 𝑚 1.5

Fixed-bed volume

𝑉𝐵𝑒𝑑 𝑚 3 3.6

Cross-sectional area

𝐴𝐵𝑒𝑑 𝑚 2 1.8

Empty Bed Contact

Time 𝑚𝑖𝑛 42

𝐸𝐵𝐶𝑇 min 42

Effective contact

time

τ 𝑚𝑖𝑛 12.5

Bed porosity ε - 0.3

Filtration Rate 𝐹𝑅 m3

m2s

0.001

Adsorption column

volume

VAds 𝑚3 5.5

Adsorption column

Height

𝐻𝐴𝑑𝑠 𝑚 3

Void-filled volume VL 𝑚3 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 130℃, 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 Ghana’s agricultural GDP and one of Ghana’s

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 in‐terms 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. Ghana’s 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

NEPAD’s, 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.

• Availability of raw material

• Cheap labour.

• Nearness of plant to target market

• Availability of electricity and water

• 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

“guesstimation” 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: 𝑈𝑆𝐷 1 = 3.84 𝐺𝐻𝐶 (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.

𝐶𝑇 = 𝐶𝐹 + 𝐶𝑊

where

𝐶𝑇 = total capital investment

𝐶𝐹 = fixed capital investment

𝐶𝑊 =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: 𝐶𝑊 = 𝐶𝐹 × 𝑓𝑊

Where 𝑓𝑊 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: 𝑓𝑊 =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;

𝐶 = 𝐶𝑂 (𝑀

𝑀𝑂)𝑛

…………………… . .8.1

Where;

C = cost of equipment

86

𝐶𝑂=cost of off-the-shelf equipment

M = capacity of equipment

𝑀𝑂= capacity of off-the-shelf equipment

𝑛= 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;

𝑀𝑎𝑟𝑠ℎ𝑎𝑙𝑙 & 𝑆𝑤𝑖𝑓𝑡 (𝑀&𝑆)𝑖𝑛𝑑𝑒𝑥 = 1100 + 20(𝑦𝑒𝑎𝑟 − 2000)……………………8.2

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 𝟑𝟐𝟐, 𝟑𝟕𝟐

Total cost of imported equipment

= starch production equipment cost + glucose preparation equipment cost

+ 𝑠𝑜𝑟𝑏𝑖𝑡𝑜𝑙 production plant equipment cost

𝑡𝑜𝑡𝑎𝑙 𝑖𝑚𝑝𝑜𝑟𝑡𝑒𝑑 𝑒𝑞𝑢𝑖𝑝𝑚𝑒𝑛𝑡 𝑐𝑜𝑠𝑡 = 135,450 + 113,957 + 72,965

= 𝑼𝑺𝑫 𝟑𝟐𝟐, 𝟑𝟕𝟐

Free On Board (FOB) = freight charges + total cost of imported equipment

But freight charges are 2% of total cost of imported equipment

𝑇𝑜𝑡𝑎𝑙 𝑐𝑜𝑠𝑡 𝑜𝑓 𝑖𝑚𝑝𝑜𝑟𝑡𝑒𝑑 𝑒𝑞𝑢𝑖𝑝𝑚𝑒𝑛𝑡 = 1.02 × 322,372

= 𝑼𝑺𝑫 𝟑𝟐𝟖, 𝟖𝟏𝟗

90

The total cost of imported processing equipment is USD 328,819 and considering the current

exchange rate of USD1= GHȻ 3.84, the Ghanaian cedi equivalent is GHȻ 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

(GHȻ)

TOTAL

COST (GHȻ)

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 GH¢ 132,200.00

Total cost of equipment = cost of imported equipment + cost of locally acquired equipment

= 1,262,667 + 132,200.00

= 𝑮𝑯¢𝟏, 𝟑𝟗𝟒, 𝟖𝟔𝟕

91

Estimates;

Spare parts = 10% of total cost of equipment

Equipment spare parts cost = 0.10 × 1,394,867

= 𝑮𝑯¢ 139,487

Handling and transportation = 0.5% of total equipment cost

= 0.05 × 1,394,867

= 𝑮𝑯¢ 𝟔𝟗, 𝟕𝟒𝟑

8.2.4 TOTAL PURCHASED EQUIPMENT COST (TPE)

Table 8.3 Total Purchase Equipment Cost

COMPONENT ESTIMATED COST (GH¢)

Equipment Cost 1,394,867

Spare parts 139487

Handling and transportation 69,743

Total TPE GH¢ 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 (GHȻ)

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) = GHȻ 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 GHȻ 1,004,807

93

Total Cost of Equipment = Direct Cost + Indirect Cost

Total Cost of Equipment = GHȻ 4,726,314

8.2.3 COST OF LAND

Cost of Land (70ft by 100ft) at Asamankese in the Eastern Region is GHȻ 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 × 3500

= 𝐆𝐇Ȼ 𝟐𝟖, 𝟎𝟎𝟎

8.3.3 ESTIMATION OF FIXED CAPITAL INVESTMENT

𝐹𝑖𝑥𝑒𝑑 𝑐𝑎𝑝𝑖𝑡𝑎𝑙 𝑖𝑛𝑣𝑒𝑠𝑡𝑚𝑒𝑛𝑡(𝐶𝐹) = 𝑡𝑜𝑡𝑎𝑙 𝑐𝑜𝑠𝑡 𝑜𝑓 𝑒𝑞𝑢𝑖𝑝𝑚𝑒𝑛𝑡 + 𝑐𝑜𝑠𝑡 𝑜𝑓 𝑙𝑎𝑛𝑑

𝐶𝐹 = 4,426,314 + 28,000

𝐆𝐇Ȼ 𝟒, 𝟕𝟓𝟒, 𝟑𝟏𝟒

8.3 ESTIMATION OF WORKING CAPITAL

The working capital 𝐶𝑤 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:

𝐶𝑊 = 𝑓𝑊 × 𝐶𝑇

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 (𝐶𝑇);

𝐶𝑇 = 𝐶𝐹 + 𝐶𝑊

𝐶𝑊 = 0.15𝐶𝑇

𝐶𝑇 = 𝐶𝐹 + 0.15𝐶𝑇

𝐶𝐹 = 𝐶𝑇 − 0.15𝐶𝑇

= 0.85𝐶𝑇

𝐶𝑇 =𝐶𝐹

0.85=

4,754,314

0.85

= 𝑮𝑯𝑪 𝟓, 𝟓𝟗𝟑, 𝟑𝟏𝟏

Hence 𝐶𝑊 = 0.15 × 5,593,311

= 𝑮𝑯𝑪 𝟖𝟑𝟖, 𝟗𝟗𝟕

SUMMARY OF INVESTMENT COST

Total Capital Investment (𝐂𝐓) GHȻ 𝟓, 𝟓𝟗𝟑, 𝟑𝟏𝟏

Working Capital (𝐂𝐖) GHȻ 838,997

Fixed Capital Investment (𝐂𝐅) GHȻ 𝟒, 𝟕𝟓𝟒, 𝟑𝟏𝟒

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(𝐂𝐌) + General Expenses (𝐆𝐄)

Manufacturing Cost(𝐂𝐌) = Direct Production Cost + Fixed Charges + Plant Overhead Cost

General Expenses (𝐆𝐄) = Administration Expenses + Distribution and Marketing Expenses

8.4.1 ESTIMATION OF MANUFACTURING COST (𝐂𝐌)

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

(GHȻ/Kg)

TOTAL (GHȻ)

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 GHȻ 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

(GHȻ)

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

= GHȻ 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 GHȻ 98,400.00

Total utilities cost = cost of power consumed + cost of water used

= 222,244 + 98,400

= 𝑮𝑯𝑪 𝟑𝟐𝟎, 𝟔𝟒𝟒

COST OF PACKAGING MATERIALS

Number of card boxes needed daily = 60

Number of required boxes annually = 18,000

Unit cost of corrugated board =GHȻ 3.00

Total cost of board annually required = GHȻ 54,000

Unit cost of low density polyethylene = GHȻ 0.80

Annual cost of low density polyethylene =GHȻ 14,400

Hence total cost of packaging material annually =GHȻ 68,400

98

Table 8.6. Estimation of Operation Labour

ITEM QUANTITY Monthly salary

(GHȻ)

TOTAL annual

salary (GHȻ)

(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 GHȻ 329,500

Social security =12.5% of total labour cost

= 0.125 × 329,500

= 𝐺𝐻𝐶 41,188

Total operating labour cost = 329,500 + 41,188

= GHȻ 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.

𝑑 =𝐶𝐹 − 𝑉𝑠

𝑛

d = annual depreciation (GHȻ/yr) =

𝐶𝐹 = the Initial Fixed Capital Investment = GHȻ 4,754,314

𝑉𝑠 = Salvage value = 10% 𝐶𝐹 = 0.10(4,754,314) = GHȻ 475,431

𝑛 = Service life or useful life of plant (yrs) = 20yrs

𝑑 =4,754,314 − 475,431

20

= 𝑮𝑯𝑪 𝟐𝟏𝟑, 𝟗𝟒𝟒

101

INSURANCE

Insurance amounts to 1% of the fixed capital investment.

𝐼𝑛𝑠𝑢𝑟𝑎𝑛𝑐𝑒 = 0.01 × 𝐶𝐹

= 0.01 ×4,754,314

= 𝐆𝐇𝐂 𝟒𝟕, 𝟓𝟒𝟑

𝐅𝐢𝐱𝐞𝐝 𝐂𝐡𝐚𝐫𝐠𝐞𝐬 = 𝐃𝐞𝐩𝐫𝐞𝐜𝐢𝐚𝐭𝐢𝐨𝐧 + 𝐈𝐧𝐬𝐮𝐫𝐚𝐧𝐜𝐞

𝑭𝒊𝒙𝒆𝒅 𝑪𝒉𝒂𝒓𝒈𝒆𝒔 = 213,944 + 47,543

= 𝑮𝑯𝑪 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 (𝐂𝐌)

COMPONENT COST FACTOR TOTAL COST(GHȻ)

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 𝐶𝐹 285,258

Laboratory charges 10% of operating labour 37,069

Total Direct Production Cost GHȻ 3,313,433

FIXED CHARGES

Depreciation 213,944

Insurance 47,543

Local Taxes 3% of 𝐶𝐹 142,629

Total Fixed Charges GHȻ 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 (𝐆𝐄)

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(𝐂𝐌) + General Expenses (𝐆𝐄)

TPC = 𝐂𝐌 + 𝐆𝐄

TPC = 3,717,549 + 0.15TPC + 0.06TPC

TPC = 3,717,549 + 0.21TPC

TPC − 0.21TPC = 3,717,549

TPC(1 − 0.21) = 3,717,549

TPC =3,717,549

0.79

𝐓𝐏𝐂 = 𝐆𝐇Ȼ 𝟒, 𝟕𝟎𝟓, 𝟕𝟓𝟖

104

Plant overhead = 0.15 × TPC

= GHȻ 705,864

General Expenses = 0.06TPC

= 0.06 × 4,705,758

= GHȻ 282,345

Total Manufacturing Cost = 3,717,549 + 0.15TPC

= 3,717,549 + 0.15 x 4,705,758

The Total Manufacturing Cost = GHȻ 4,423,413

SUMMARY OF TOTAL PRODUCTION COST

DIRECT PRODUCTION COST GHȻ 3,313,433

FIXED CHARGES GHȻ 404,116

PLANT OVERHEAD COST GHȻ 705,864

GENERAL EXPENSES GHȻ 282,345

MANUFACTURING COST (𝐂𝐌) GHȻ 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

𝑆𝑒𝑙𝑙𝑖𝑛𝑔 𝑃𝑟𝑖𝑐𝑒 = 𝑇𝑜𝑡𝑎𝑙 𝐶𝑜𝑠𝑡 × (1 + 𝑚𝑎𝑟𝑘 − 𝑢𝑝 𝑝𝑒𝑟𝑐𝑒𝑛𝑡)

Production rate of sorbitol = 3,000 kg per day

106

Total annual production = 3,000𝑘𝑔

𝑑𝑎𝑦 × 300

𝑑𝑎𝑦

𝑦𝑒𝑎𝑟

= 900,000 𝑘𝑔

Therefore, the total annual production is 900,000 𝑘𝑔

Number of 50 kg bags per year = 900,000

50

= 18,000 𝑏𝑜𝑥𝑒𝑠

Total Production Cost (TPC) = GHȻ 4,705,758

The cost of producing a 50 kg box =𝑇𝑃𝐶

Number of 50 kg bags per year

= 4,705,758

18,000

= 𝐺𝐻𝐶 261.

Using a total mark-up percent of 65 %

𝑆𝑒𝑙𝑙𝑖𝑛𝑔 𝑃𝑟𝑖𝑐𝑒 = 𝑇𝑜𝑡𝑎𝑙 𝐶𝑜𝑠𝑡 × (1 + 𝑚𝑎𝑟𝑘 − 𝑢𝑝 𝑝𝑒𝑟𝑐𝑒𝑛𝑡)

= 261 × (1 + 0.65)

= GHȻ 431

Therefore, a 50 Kg box of sorbitol will be sold for GHȻ430

Revenue at the end of production year = 𝑆𝑒𝑙𝑙𝑖𝑛𝑔 𝑃𝑟𝑖𝑐𝑒 × No. of 50 kg bags produced annually

= 430 × 18,000

= 𝐺𝐻𝐶 7,740,000

Hence, the Revenue at the end of production year is 𝐺𝐻𝐶 7,740,000

107

DATA

Fixed Capital Investment (CF) = GHȻ 4,754,314

Working Capital (Cw) =GHȻ 838,997

Total Capital Investment (CT) = GHȻ 5,593,311

Total Production Cost (TPC) = GHȻ 4,705,758

Manufacturing Cost (CM) = GHȻ 4,423,413

Salvage Value (VS) = GHȻ 475,431

Depreciation (D) =GHȻ 213,944

ECONOMIC ENVIRONMENT

Depreciation coefficient 𝒅𝒄 =𝟏

𝒏=

1

20= 0.05

Depreciation d = GH Ȼ 213,944 per year

Interest rate, i (discount rate) = 35 %

Tax rate, t =12.50%

Plant life, n= 20 years

Capital recovery, 𝑒 =𝑖(1+𝑖)𝑛

(1+𝑖)𝑛−1

= 0.35(1+0.35)20

(1+0.35)20−1 =0.35

108

8.5.2 GROSS PROFIT (PG)

𝐺𝑟𝑜𝑠𝑠 𝑃𝑟𝑜𝑓𝑖𝑡 = 𝑅𝑒𝑣𝑒𝑛𝑢𝑒 − 𝐶𝑜𝑠𝑡 𝑜𝑓 𝑀𝑎𝑛𝑢𝑓𝑎𝑐𝑡𝑢𝑟𝑖𝑛𝑔

PG = 7,740,000 −4,423,413

PG = GH Ȼ 3,316,587

8.5.3 TAXABLE INCOME (R)

𝑇𝑎𝑥𝑎𝑏𝑙𝑒 𝐼𝑛𝑐𝑜𝑚𝑒 = Gross Profit − 𝑑𝑒𝑝𝑟𝑒𝑐𝑖𝑎𝑡𝑖𝑜𝑛 𝑐𝑜𝑒𝑓𝑓𝑖𝑐𝑖𝑒𝑛𝑡 × Fixed Capital Investment

𝑅 = 𝑃𝐺 − 𝒅𝒄 × 𝐶𝐹

= 3,316,587 − 0.05 × 4,754,314

= 𝐺𝐻Ȼ 3,078,871

8.5.4 INCOME TAX (T)

𝐼𝑛𝑐𝑜𝑚𝑒 𝑇𝑎𝑥 = Taxable Income × Tax rate

𝑇 = 𝑅 × 𝑡

= 3,078,871 × 0.125

= 𝐺𝐻Ȼ 348,859

8.5.5 ANNUAL PROFIT AFTER TAX (P)

𝐴𝑛𝑛𝑢𝑎𝑙 𝑃𝑟𝑜𝑓𝑖𝑡 𝑎𝑓𝑡𝑒𝑟 𝑇𝑎𝑥 = 𝐺𝑟𝑜𝑠𝑠 𝑃𝑟𝑜𝑓𝑖𝑡 − 𝐼𝑛𝑐𝑜𝑚𝑒 𝑇𝑎𝑥

P = PG – T

= 3,316,587 − 348,859

= GH Ȼ 2,931,728

109

8.5.6 ANNUAL CASH FLOW (CF)

Annual Cash Flow = Annual Profit after Tax + Depreciation

𝐶𝐹 = 𝑃 + 𝑑

= 2,931,728+ 213,944

= GH Ȼ 3,145,672

8.5.7 NET PROFIT (PN)

Net Profit = Annual Cash Flow + Recovery factor x Fixed Capital Investment

PN= CF - e𝐶𝐹

= 3,145,672– 0.35 × 4,754,314

= GH Ȼ 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=𝑇𝑜𝑡𝑎𝑙 𝐶𝑎𝑝𝑖𝑡𝑎𝑙 𝐼𝑛𝑣𝑒𝑠𝑡𝑚𝑒𝑛𝑡

𝐴𝑛𝑛𝑢𝑎𝑙 𝑃𝑟𝑜𝑓𝑖𝑡 𝑎𝑓𝑡𝑒𝑟 𝑇𝑎𝑥

𝑃𝐵𝑃 =𝐶𝑇

𝑃

=5,593,311

2,931,728 = 1.9 𝑦𝑒𝑎𝑟𝑠

8.6.2 DISCOUNTED PAYBACK PERIOD

DPB =

[ ln (

11 − i(PBP)

)

ln (1 + i)

]

=

[ ln (

11 − 0.35(1.9)

)

ln (1 + 0.35)

]

𝐃𝐏𝐁 = 𝟑. 𝟔 𝐲𝐞𝐚𝐫𝐬

111

8.6.3 CUMULATIVE CASH FLOW

CCF = −CT + (n x CF)

CCF = −5,593,311 + (20 x 3,145,672)

CCF = GHȻ 57,320,129

8.7 SENSITIVITY ANALYSIS

Return on Investment (ROI):

It is expressed on annual percentage basis:

𝑅𝑒𝑡𝑢𝑟𝑛 𝑜𝑛 𝑖𝑛𝑣𝑒𝑠𝑡𝑚𝑒𝑛𝑡 = 𝑃

𝐶𝑇× 100%

ROI = 2,931,728

5,593,311 × 100%

= 52.4%

Therefore, the Return on Investment (ROI) is 52.4%

Net Present Value (NPV)

Capital Recovery Factor, e = 0.35

NPV = [−CT +P

e]

NPV = [−5,593,311 +2,931,728

0.35]

NPV = GHȻ 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

REFERENCE

1) Afoakwa, E. et al., (2012). Chemical composition and cyanogenic potential of traditional

and high yielding CMD resistant cassava (Manihot esculenta ), International Food

Research Journal 19(1), 175-181.

2) Aiyer, P. V, Arts, S.P.T. & College, S., (2005). Amylases and their applications. ,

4(December), pp.1525–1529.

3) Breuninger W.F., Kuakoon P, and Klanarong S, (2009). Tapioca/Cassava Starch:

Production and Use. Chemistry and Technology, Third Edition Pp 541-568. ISBN: 978-

0-12-746275-2.

4) Charles, A.L., Sriroth, K. and Huang, T.C. (2005). Proximate composition, mineral

contents, hydrogen cyanide and phytic acid of 5 cassava genotypes. Food Chemistry 92:

615–20.

5) Chilton CH, (1960). Cost Engineering in the Process Industries. McGraw-Hill, New

York.

6) Dziedzoave, N.T., Abass, A. B., Amoa‐Awua, W.K.A. and Sablah, M. (2006). Quality

Management manual for production of high quality cassava flour. (Adegoke, G.O. and

Brimer, L. eds). International Institute of Tropical Agriculture (IITA).

7) Eckhard, W., (2012). Adsorption Technology in Water Treatment, Berlin/Boston: Walter

de Gruyter GmbH & Co.

8) Fellows P.J., (2000). Food Processing Technology Principles and Practice. 2nd Edition.

Published by Woodhead Publishing Limited Abington Hall, Abington Cambridge CB1

6AH, England. pp 323-327.

116

9) Fontana, J.D. et al., (2008). Starch Depolymerization with Diluted Phosphoric Acid and

Application of the Hydrolysate in Astaxanthin Fermentation. , 46(3), pp.305–310.

10) Guillaume,D. Dominique Dufour, Claude Marouzé, Mai Le Thanhb, Pierre-André

Maréchalc,(2008). Cassava Starch Processing at Small Scale in North Vietnam.

Starch/Stärke 60 (2008) 358–372.

11) Heinen A. W., Peters J. A., and H. Van Bekkum. P, “Hydrogenation of fructose on Ru/C

catalysts,” Carbohydrate Research, vol. 328, no. 4, pp. 449-457, Oct. 2000.

12) Hillocks, R.J. (2002). Cassava in Africa. In: Hillocks, R.J., Thresh, J.M. and Bellotti,

A.C. (eds) Cassava: Biology, Production and Utilization. CAB International, pp 41 – 54.

13) Hobbs, L., (2009). Sweeteners from Starch : Production , Properties and Uses. ISBN

9780127462752.

14) Kassim Bin A , Rice C.L (1980).Formation of sorbitol by catholytic reduction of glucose.

Journal of Applied electrochemistry.11(261-267).

15) Klanarong, S, Kuakoon, P, Wanlapatitc, S., Christopher G., (2000). Cassava Starch

Technology: The Thai Experience. Starch/Stärke 52 (20) 439–449.

16) Lebot, V. (2009). Cassava: postharvest quality and marketing. Pp. 413 in V. Lebot, ed.

Tropical root and tuber crops cassava, sweet potato, yams and aroids. Crop Production

Science in Horticulture No. 17, CABI, Wallingford, England.

17) Montagnac, J.A., Davis, C.R. and Tanumihardjo, S.A. (2009). Nutritional Value of

Cassava for use as a Staple Food and Recent Advances for Improvement. Comprehensive

Review in Food Science and Food Safety 8: 181-188.

117

18) Meridian Institute, Innovations for Agricultural Value Chains in Africa: Applying

Science and Technology to Enhance Cassava, Dairy, and Maize Value Chains. Cassava

Value Chain Overview, 2009.

19) Padonou, W., Mestres, C. and Nago, M.C.,(2005). The quality of boiled cassava roots:

instrumental characterization and relationship with physicochemical properties and

sensorial properties. Food Chemistry 89: 261–270.

20) Park,K, Pintauro,P.N., Baizer,M.M. (1985), Journal of Electochemistry. Soc. 132 1850

21) Peters M.S and Timmerhaus K.D. (1991). Plant Design and Economics for Chemical

Engineers. McGraw- Hill Inc. New York.

22) Sanjust, E. et al.(2004). Xylose production from durum wheat bran: enzymic versus

chemical methods. Food Science and Technology International, v. 10, n. 1, p. 11-14.

23) Saravacos G.D, Kostaropoulos A.E, (2002). Handbook of Food Processing Equipment.

Kluwer Academic / Plenum Publ, New York.

24) Schwartz, D. & Whistler, R.L., (2009). History and Future of Starch Third Edit., Elsevier

Inc. Available at: http://dx.doi.org/10.1016/B978-0-12-746275-2.00001-X

25) Silveira, M.M., Sales, R., Lemmel, C., Jonas, R.(1995). Glucose–fructose activity in six

strains of Zymomonas mobilis. Arq. Biol. Tecnol. 38, 619–622.

26) Wahyuningtyas, A., Roto, R. & Kuncaka, A.,(2015).. Aian Journal of Chemistry, 28(5),

pp.987–992.

27) Wisbeck, E., Silveira, M.M., Ninow, J., Jonas, R.(1997).Evaluation of the flocculent

strain Zymomonas mobilis Z1-81 for the production of sorbitol and gluconic acid. J.Basic

Microbiol. 37, 445–449.

118

28) Wisniak J. and Simon,R., Hydrogenation of glucose, fructose, and their mixtures.

Industrial & Engineering Chemistry Product Research and Development, vol. 18, no. 1,

pp. 50–57, 1979.

29) Zacharious, M., Scopes, R.K., (1986). Glucose–fructose oxidoreductase, a new enzyme

isolated from Zymomonas mobilis that is responsible for sorbitol production. J.

Bacteriol.167, 803–809.

30) Zvinavashe, E., Elbersen, H. W., Slingerland, M., Kolijn, S. and Sanders, J. (2011).

Cassava for food and energy: exploring potential benefits of processing of cassava into

cassava flour and bioenergy at farmstead and community levels in rural Mozambique.

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

𝐴𝑜𝑢𝑛𝑡 𝑜𝑓 𝑤𝑎𝑠𝑡𝑒 𝑟𝑒𝑚𝑜𝑣𝑒𝑑, 𝐴𝑜1 =1

100× 𝐴𝑖1 = 0.01 × 15000 = 150𝐾𝑔/𝑑𝑎𝑦

𝑅𝑜𝑜𝑡𝑠 𝐴𝑓𝑡𝑒𝑟 𝑆𝑜𝑟𝑡𝑖𝑛𝑔, 𝐴𝑜2 = 𝐶𝑎𝑠𝑠𝑎𝑣𝑎 𝑟𝑒𝑐𝑖𝑒𝑣𝑒𝑑 − 𝑊𝑎𝑠𝑡𝑒 𝑟𝑒𝑚𝑜𝑣𝑒𝑑

= 15000 − 150 = 14850𝐾𝑔/𝑑𝑎𝑦

Washing and Peeling

Amount of water used for washing, Bi1 =141.4

100× 𝐴𝑜1 = 1.414 × 14850

= 20997.9𝐾𝑔/𝐷𝐴𝑌

𝐴𝑚𝑜𝑢𝑛𝑡 𝑜𝑓 𝑤𝑎𝑠𝑡𝑒 𝑤𝑎𝑡𝑒𝑟, 𝐵𝑜2 =99.641

100× 𝐵𝑖1 = 0.99641 ×

20997.9𝐾𝑔

𝑑𝑎𝑦

= 20922.52𝐾𝑔/𝑑𝑎𝑦

𝐴𝑚𝑜𝑢𝑛𝑡 𝑜𝑓 𝑝𝑒𝑒𝑙𝑠, 𝐵𝑜1 =1.5

100× 𝐴𝑜2 = 0.015 × 14850

=222.75𝐾𝑔

𝑑𝑎𝑦

𝑃𝑒𝑒𝑙𝑒𝑑 𝑅𝑜𝑜𝑡𝑠, 𝐵𝑜2 = 𝐴02 + 𝐵𝑖1 − (𝐵𝑜1 + 𝐵𝑜2)

= 14850 + 20997.9 − (20997.9 + 222.75)

= 14702.63𝐾𝑔

𝑑𝑎𝑦

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

𝑤𝑎𝑡𝑒𝑟 𝑎𝑑𝑑𝑒𝑑 𝑡𝑜 𝑟𝑎𝑠𝑝𝑒𝑟, 𝐶𝑖2 =36.5

100× 𝐶𝑖1

𝑏𝑢𝑡 𝐶𝑖1 = 𝐵𝑜2 =14702.63𝐾𝑔

𝑑𝑎𝑦

𝐶𝑖2 = 0.365 × 14702.63 𝑘𝑔/𝑑𝑎𝑦

=5366.46𝑘𝑔

𝑑𝑎𝑦

𝑅𝑎𝑠𝑝𝑒𝑑 𝑝𝑢𝑙𝑝, 𝐶𝑜1 = 𝐶𝑖1 + 𝐶𝑖2

= 14702.63 + 5366.46 = 20069.09𝐾𝑔

𝑑𝑎𝑦

1st Extraction

𝑠𝑢𝑙𝑝ℎ𝑢𝑟 𝑤𝑎𝑡𝑒𝑟 𝑎𝑑𝑑𝑒𝑑, 𝐷𝑖1 =32

100× 𝐶𝑖1 = 0.32 ×

14702.63 𝐾𝑔

𝑑𝑎𝑦

𝐷𝑖1 =6422.11𝐾𝑔

𝑑𝑎𝑦

𝑊𝑎𝑡𝑒𝑟 𝑢𝑠𝑒𝑑, 𝐷𝑖2 =40

100× 𝐶𝑖1 = 0.40 ×

14702.63𝐾𝑔

𝑑𝑎𝑦

𝐷𝑖2 =5881.05𝐾𝑔

𝑑𝑎𝑦

𝑐𝑜𝑎𝑟𝑠𝑒 𝑠𝑡𝑎𝑟𝑐ℎ 𝑟𝑒𝑐𝑦𝑐𝑙𝑒𝑑, 𝐷𝑜1 =35

100× (𝐶𝑖1 + 𝐷𝑖1 + 𝐷𝑖2)

= 0.35 × (14702.63 + 6422.11 + 5881.05)

123

= 0.35 ×27005.79𝐾𝑔

𝑑𝑎𝑦

=11330.29𝐾𝑔

𝑑𝑎𝑦

𝐹𝑖𝑛𝑒 𝑆𝑙𝑢𝑟𝑟𝑦, 𝐷𝑜2 = 𝐷𝑖1 + 𝐷𝑖2 + 𝐷𝑜1

= 6422.11 + 5881.05 + 11330.29

= 21041.97𝐾𝑔

𝑑𝑎𝑦

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

𝑾𝒂𝒕𝒆𝒓 𝒖𝒔𝒆𝒅,𝑯𝒊𝟐 =45

100× 𝐻𝑖1

But 𝐻𝑖1 = 𝐷𝑜1 = 11330.29𝐾𝑔

𝑑𝑎𝑦

Hence 𝐻𝑖2 = 0.45 × 11330.29𝐾𝑔

𝑑𝑎𝑦

= 5098.63 𝐾𝑔/𝑑𝑎𝑦

𝑭𝒊𝒏𝒆 𝒔𝒍𝒖𝒓𝒓𝒚 𝒆𝒙𝒕𝒓𝒂𝒄𝒕𝒆𝒅,𝐻𝑜1 = 65

100× 𝐻𝑖1 = 0.65 × 11330.29

= 7364.69𝐾𝑔

𝑑𝑎𝑦

125

𝑪𝒂𝒔𝒔𝒂𝒗𝒂 𝑷𝒖𝒍𝒑,𝐻𝑜2 = 𝐻𝑖1 + 𝐻𝑖2 − 𝐻𝑜1

= 11330.29 + 5098.63 − 7364.69 = 9064.23𝐾𝑔

𝑑𝑎𝑦

PULP PRESS

𝑠𝑡𝑎𝑟𝑐ℎ 𝑠𝑙𝑢𝑟𝑟𝑦 𝑜𝑏𝑡𝑎𝑖𝑛𝑒𝑑, 𝐼𝑜1 =31

100× 𝐻𝑜2

= 0.31 × 9064.23 = 2809.91𝐾𝑔

𝑑𝑎𝑦

𝑠𝑡𝑎𝑟𝑐ℎ 𝐶𝑎𝑘𝑒, 𝐼𝑜2 = 𝐻𝑜2 − 𝐼𝑜1

= 9064.23 − 2809.91 = 6254.32𝐾𝑔

𝑑𝑎𝑦

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

𝑤𝑎𝑡𝑒𝑟 𝑖𝑛𝑝𝑢𝑡, 𝐸𝑖2 =30

100× (𝐸𝑖1 + 𝐻𝑜1)

= 0.30 × (21041.97 + 7364.69)

= 4261𝐾𝑔

𝑑𝑎𝑦

𝑠𝑢𝑙𝑝ℎ𝑢𝑟 𝑤𝑎𝑡𝑒𝑟 𝑖𝑛𝑝𝑢𝑡, 𝐸𝑖3 =15

100× 𝐸𝑖1

= 0.15 × 28406.65

= 8522𝐾𝑔

𝑑𝑎𝑦

127

𝑤𝑎𝑠𝑡𝑒𝑤𝑎𝑡𝑒𝑟, 𝐸𝑜1 =67

100× (𝐸𝑖1 + 𝐻𝑜1 + 𝐸𝑖2 + 𝐸𝑖3)

= 0.67 × (21041.97 + 7364.69 + 4261 + 8522)

= 27597.07𝐾𝑔

𝑑𝑎𝑦

𝑐𝑜𝑛𝑐𝑒𝑛𝑡𝑟𝑎𝑡𝑒𝑑 𝑠𝑙𝑢𝑟𝑟𝑦, 𝐸𝑜2 = 𝐸𝑖1 + 𝐻𝑜1 + 𝐸𝑖2 + 𝐸𝑖3 − 𝐸𝑜1

= 21041.97 + 7364.69 + 4261 + 8522 − 27597.07

= 13592.58𝐾𝑔

𝑑𝑎𝑦

Dewatering

𝑤𝑎𝑠𝑡𝑒𝑤𝑎𝑡𝑒𝑟, 𝐹𝑜1 =67

100× 𝐸𝑜2

= 0.67 × 13592.58

= 9107.03𝐾𝑔

𝑑𝑎𝑦

𝑠𝑡𝑎𝑟𝑐ℎ 𝑐𝑘𝑎𝑒, 𝐹𝑜2 = 𝐸𝑜2 − 𝐹𝑜1

= 13592.58 − 9107.03

= 4485.55𝐾𝑔

𝑑𝑎𝑦

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

𝑤𝑙 = 0.2 × 6048.8 = 1209.8 kg/day

𝑆𝑖 + 𝐸1 = 𝑆𝑜1 + 𝑤𝑙

6048.8 + 12 = 𝑆𝑜1 + 1209.8

= 4851 𝑘𝑔/𝑑𝑎𝑦

partially hydrolyzed starch, 𝑆𝑜1 = 4851 𝑘𝑔/𝑑𝑎𝑦

The quantity of partially hydrolyzed starch from the liquefaction process is 4851 𝑘𝑔/𝑑𝑎𝑦

130

SACCHARIFICATION

ASSUMPTON

The water loss during the liquefaction process is negligible.

𝑆01 + 𝐸2 =, 𝐺01

𝐺𝑜1 = 4851 + 9.7

𝐺𝑜1 = 4860.7 𝑘𝑔/𝑑𝑎𝑦

Crude Glucose, 𝐺𝑜1 = 4860.7 𝑘𝑔/𝑑𝑎

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

𝑆𝑎 = 243 𝑘𝑔/𝑑𝑎𝑦

Total mass in =Total mass out

𝐺𝑂1 + 𝑆𝐴 = 𝐺02

𝐺02 = 4861 + 243

𝐺02 = 50103.7

Acidified Glucose, 𝐺02 = 5103.7 𝑘𝑔/𝑑𝑎𝑦

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% × 𝐺𝑜2

= 1% × 5104

= 51𝑘𝑔/𝑑𝑎𝑦

Mass in =Mass out

𝐺02 = 𝐺03 + 𝑅01

133

𝐺𝑜3 = 5104 − 51

𝐺𝑜3 = 5053

Refined Glucose syrup, 𝐺𝑜3 = 5053 𝑘𝑔/𝑑𝑎𝑦

The quantity of the refined Glucose syrup after the adsorption is 5053 𝑘𝑔/𝑑𝑎𝑦

ION EXCHANGE (CATION)

ASSUMPTION

The residues, Ro2 is 0.3% of the incoming glucose syrup=0.3% × 𝐺𝑜3

= 0.3% × 5104

= 15.2 𝑘𝑔/𝑑𝑎𝑦

134

CALCULATIONS

𝑀𝑎𝑠𝑠 𝐼𝑛 = 𝑀𝑎𝑠𝑠 𝑂𝑢𝑡

𝐺𝑜3 = 𝐺𝑜4 + 𝑅02

𝐺𝑜4 = 5053 − 15.2

Refined Glucose syrup, 𝐺𝑜4 = 5037.8 𝑘𝑔/𝑑𝑎𝑦

The quantity of the De-ionised Glucose syrup after the adsorption is 5038 𝑘𝑔/𝑑𝑎𝑦

ION EXCHANGE (ANION)

ASSUMPTION

The residues,Ro2 is 0.3% of the incoming glucose syrup=0.3% × 𝐺𝑜3

= 0.3% × 5104

= 15.2 𝑘𝑔/𝑑𝑎𝑦

135

CALCULATIONS

Mass in = Mass out

𝐺𝑜4 = 𝑅03 + 𝐺𝑜5

𝐺𝑜5 = 𝐺𝑜4 − 𝑅𝑜5

𝐺𝑜5 = 5037.8 − 15

final glucose, 𝐺𝑜5 = 5022.8 𝑘𝑔/𝑑𝑎𝑦

The quantity of the Refined Glucose syrup after the adsorption is 5023 𝑘𝑔/𝑑𝑎𝑦

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

𝑛 (𝐶6𝐻12𝑂6)

𝑛(𝐻2)= 1

𝐼𝑛𝑖𝑡𝑖𝑎𝑙 𝑎𝑚𝑜𝑢𝑛𝑡 𝑜𝑓 𝑑𝑒𝑥𝑡𝑟𝑜𝑠𝑒 = 𝐷𝑒𝑥𝑡𝑟𝑜𝑠𝑒 𝑓𝑙𝑜𝑤 𝑟𝑎𝑡𝑒 ×1

𝑚𝑜𝑙𝑎𝑟 𝑚𝑎𝑠𝑠 𝑜𝑓 𝑑𝑒𝑥𝑡𝑟𝑜𝑠𝑒

= 5503 ×1

180.1559

= 30.5 𝑚𝑜𝑙 𝑜𝑓 𝑑𝑒𝑥𝑡𝑟𝑜𝑠𝑒 𝑠𝑦𝑟𝑢𝑝.

𝐼𝑛𝑖𝑡𝑖𝑎𝑙 𝑎𝑚𝑜𝑢𝑛𝑡 𝑜𝑓 ℎ𝑦𝑑𝑟𝑜𝑔𝑒𝑛 = 𝐻𝑦𝑑𝑟𝑜𝑔𝑒𝑛 𝑓𝑙𝑜𝑤 𝑟𝑎𝑡𝑒 ×1

𝑚𝑜𝑙𝑎𝑟 𝑚𝑎𝑠𝑠 𝑜𝑓 ℎ𝑦𝑑𝑟𝑜𝑔𝑒𝑛

= 1005 ×1

2.016

= 498.3 𝑚𝑜𝑙 𝑜𝑓 ℎ𝑦𝑑𝑟𝑜𝑔𝑒𝑛 𝑔𝑎𝑠

𝑅𝑎𝑡𝑖𝑜 𝑜𝑓 𝑡ℎ𝑒 𝑖𝑛𝑖𝑡𝑖𝑎𝑙 𝑎𝑚𝑜𝑢𝑛𝑡𝑠 𝑜𝑓 𝑟𝑒𝑎𝑐𝑡𝑎𝑛𝑡 = 𝑛 (𝐶6𝐻12𝑂6)

𝑛(𝐻2)

= 30.5

498.3= 0.061

𝑆𝑖𝑛𝑐𝑒 𝑡ℎ𝑒 𝑟𝑎𝑡𝑖𝑜 𝑜𝑓 𝑡ℎ𝑒 𝑖𝑛𝑖𝑡𝑖𝑎𝑙 𝑎𝑚𝑜𝑢𝑛𝑡 𝑜𝑓 𝑟𝑒𝑎𝑐𝑡𝑎𝑛𝑡 𝑖𝑠 𝑙𝑒𝑠𝑠 𝑡ℎ𝑎𝑛 𝑡ℎ𝑒 𝑠𝑡𝑜𝑖𝑐ℎ𝑖𝑜𝑚𝑒𝑡𝑟𝑖𝑐 𝑟𝑎𝑡𝑖𝑜, 𝑖𝑡 𝑖𝑚𝑝𝑙𝑖𝑒𝑠 𝑑𝑒𝑥𝑡𝑟𝑜𝑠𝑒 𝑖𝑠

𝑡ℎ𝑒 𝑙𝑖𝑚𝑖𝑡𝑖𝑛𝑔 𝑟𝑒𝑎𝑐𝑡𝑎𝑛𝑡. (𝑇ℎ𝑒 𝑎𝑚𝑜𝑢𝑛𝑡 𝑜𝑓 𝑑𝑒𝑥𝑡𝑟𝑜𝑠𝑒 𝑟𝑒𝑎𝑐𝑡𝑖𝑛𝑔 𝑤𝑖𝑙𝑙 𝑑𝑒𝑡𝑒𝑟𝑚𝑖𝑛𝑒 𝑡ℎ𝑒 𝑎𝑚𝑜𝑢𝑛𝑡 𝑜𝑓 𝑠𝑜𝑟𝑏𝑖𝑡𝑜𝑙 )

138

𝑇ℎ𝑒 𝑡ℎ𝑒𝑜𝑟𝑒𝑡𝑖𝑐𝑎𝑙 𝑦𝑖𝑒𝑙𝑑 𝑜𝑓 𝑠𝑜𝑟𝑏𝑖𝑡𝑜𝑙 𝑖𝑛 𝑚𝑜𝑙𝑒𝑠 = 𝑚𝑜𝑙𝑒𝑠 𝑜𝑓 𝑑𝑒𝑥𝑡𝑟𝑜𝑠𝑒 ×1 𝑚𝑜𝑙 𝑜𝑓 𝑠𝑜𝑟𝑏𝑖𝑡𝑜𝑙

1 𝑚𝑜𝑙 𝑜𝑓 𝑑𝑒𝑥𝑡𝑟𝑜𝑠𝑒

= 30.5 𝑚𝑜𝑙 𝑜𝑓 𝑠𝑜𝑟𝑏𝑖𝑡𝑜𝑙.

𝑇ℎ𝑒 𝑚𝑎𝑠𝑠 𝑜𝑓 𝑠𝑜𝑟𝑏𝑖𝑡𝑜𝑙 𝑝𝑟𝑜𝑑𝑢𝑐𝑒𝑑 = 𝑚𝑜𝑙𝑒𝑠 × 𝑀𝑜𝑙𝑎𝑟 𝑚𝑎𝑠𝑠

= 30.5 × 182.17

𝑚𝑎𝑠𝑠 𝑜𝑓 𝑠𝑜𝑟𝑏𝑖𝑡𝑜𝑙 𝑝𝑟𝑜𝑑𝑢𝑐𝑒𝑑 𝑖𝑛 𝑎 𝑑𝑎𝑦 𝑖𝑠 5474𝑘𝑔

𝑑𝑎𝑦

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

𝐵 = 𝐶 + 𝐷

B =Flow rate of sorbitol solution =5474 kg/day

C = Flow rate of water evaporation

D = Flow rate of concentrated sorbitol solution

5474𝑘𝑔

𝑑𝑎𝑦= 𝐶 + 𝐷 …. (1)

140

Component mass balance on evaporation

Sorbitol content (S)

Component mass in = Component mass out

𝑀𝑎𝑠𝑠 𝑜𝑓 𝑠𝑜𝑟𝑏𝑖𝑡𝑜𝑙 = % 𝑠𝑜𝑟𝑏𝑖𝑡𝑜𝑙 × 𝑓𝑙𝑜𝑤 𝑟𝑎𝑡𝑒

𝐵(𝑆𝐵) = 𝐶(𝑆𝐶) + 𝐷(𝑆𝐷)

𝑆𝐵 = % 𝑆𝑜𝑟𝑏𝑖𝑡𝑜𝑙 𝑜𝑓 𝑠𝑜𝑟𝑏𝑖𝑡𝑜𝑙 𝑠𝑜𝑙𝑢𝑡𝑖𝑜𝑛 = 55%

𝑆𝐷 = % 𝑆𝑜𝑟𝑏𝑖𝑡𝑜𝑙 𝑜𝑓 𝑐𝑜𝑛𝑐𝑒𝑛𝑡𝑟𝑎𝑡𝑒𝑑 𝑠𝑜𝑟𝑏𝑖𝑡𝑜𝑙 = 70%

5474(0.55) = 𝐷(0.70)

𝐵𝑢𝑡 𝑓𝑟𝑜𝑚 (1), 𝐷 =3010.7

0.7

𝐷 = 4301𝑘𝑔

𝑑𝑎𝑦

𝐶 = 1173𝑘𝑔

𝑑𝑎𝑦

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

𝐷 = 𝐸 + 𝐹 . . (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)

𝑀𝑎𝑠𝑠 𝑜𝑓 𝑠𝑜𝑟𝑏𝑖𝑡𝑜𝑙 = % 𝑠𝑜𝑟𝑏𝑖𝑡𝑜𝑙 × 𝐹𝑙𝑜𝑤 𝑟𝑎𝑡𝑒

𝐷(𝑆𝐷) = 𝐸(𝑆𝐸) + 𝐹(𝑆𝐹)

𝑆𝐷 = % 𝑠𝑜𝑟𝑏𝑖𝑡𝑜𝑙 𝑐𝑜𝑛𝑡𝑒𝑛𝑡 𝑜𝑓 𝑐𝑜𝑛𝑐𝑒𝑛𝑡𝑟𝑎𝑡𝑒𝑑 𝑠𝑜𝑟𝑏𝑖𝑡𝑜𝑙 = 70%

𝑆𝐹 = % 𝑠𝑜𝑟𝑏𝑖𝑡𝑜𝑙 𝑜𝑓 𝑐𝑟𝑦𝑠𝑡𝑎𝑙 𝑠𝑜𝑟𝑏𝑖𝑡𝑜𝑙 = 90%

4301(0.7) = 0.90𝐹

3010.7 = 0.90𝐹

𝐹 = 3345 𝑘𝑔

𝑑𝑎𝑦

𝐹 =956 𝑘𝑔

𝑑𝑎𝑦

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

𝐹 = 𝐺 + 𝐻 ……… . . (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)

𝑀𝑎𝑠𝑠 𝑜𝑓 𝑆𝑜𝑟𝑏𝑖𝑡𝑜𝑙 = % 𝑠𝑜𝑟𝑏𝑖𝑡𝑜𝑙 × 𝐹𝑙𝑜𝑤 𝑟𝑎𝑡𝑒

𝐹(𝑆𝐹) = 𝐺(𝑆𝐺) + 𝐻(𝑆𝐻)

𝑆𝐹 = % 𝑠𝑜𝑟𝑏𝑖𝑡𝑜𝑙 𝑜𝑓 𝑐𝑟𝑦𝑠𝑡𝑎𝑙 𝑠𝑜𝑟𝑏𝑖𝑡𝑜𝑙 = 90%

𝑆𝐹𝐻 = % 𝑠𝑜𝑟𝑏𝑖𝑡𝑜𝑙 𝑜𝑓 𝑠𝑜𝑟𝑏𝑖𝑡𝑜𝑙 𝑐𝑎𝑘𝑒 = 95%

3345(0.90) = 0.95 𝐻

𝐻 =3169 𝑘𝑔

𝑑𝑎𝑦

𝐺 =176 𝑘𝑔

𝑑𝑎𝑦

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

𝑇𝑜𝑡𝑎𝑙 𝑚𝑎𝑠𝑠 𝑖𝑛 = 𝑡𝑜𝑡𝑎𝑙 𝑚𝑎𝑠𝑠 𝑜𝑢𝑡

𝐻 = 𝐼 + 𝐽

𝑊ℎ𝑒𝑟𝑒, 𝐻 = 𝑡ℎ𝑒 𝑚𝑎𝑠𝑠 𝑓𝑙𝑜𝑤 𝑟𝑎𝑡𝑒 𝑜𝑓 𝑚𝑜𝑖𝑠𝑡 𝑠𝑜𝑟𝑏𝑖𝑡𝑜𝑙 𝑐𝑎𝑘𝑒𝑠

𝐼 = 𝐹𝑙𝑜𝑤 𝑟𝑎𝑡𝑒 𝑜𝑓 𝑚𝑜𝑖𝑠𝑡𝑢𝑟𝑒 𝑟𝑒𝑚𝑜𝑣𝑒𝑑 𝑑𝑢𝑟𝑖𝑛𝑔 𝑑𝑟𝑦𝑖𝑛𝑔

𝐽 = 𝐹𝑙𝑜𝑤 𝑟𝑎𝑡𝑒 𝑜𝑓 𝑑𝑟𝑖𝑒𝑑 𝑠𝑜𝑟𝑏𝑖𝑡𝑜𝑙 𝑐𝑟𝑦𝑠𝑡𝑎𝑙𝑠

Component mass balance

Component mass in = component mass out

146

Moisture content (M)

𝑀𝑎𝑠𝑠 𝑜𝑓 𝑠𝑜𝑟𝑏𝑖𝑡𝑜𝑙 = % 𝑠𝑜𝑟𝑏𝑖𝑡𝑜𝑙 × 𝑓𝑙𝑜𝑤 𝑟𝑎𝑡𝑒

𝑆𝐻 = % 𝑠𝑜𝑟𝑏𝑖𝑡𝑜𝑙 𝑜𝑓 𝑠𝑜𝑟𝑏𝑖𝑡𝑜𝑙 𝑐𝑎𝑘𝑒 = 95%

𝑆𝐽 = % 𝑠𝑜𝑟𝑏𝑖𝑡𝑜𝑙 𝑜𝑓 𝑑𝑟𝑖𝑒𝑑 𝑠𝑜𝑟𝑏𝑖𝑡𝑜𝑙 𝑐𝑟𝑦𝑠𝑡𝑎𝑙 = 99%

𝐻(𝑆𝐻) = 𝐼(𝑆𝐼) + 𝐽(𝑆𝐽)

3169(0.95) = 0.99𝐽

𝐽 = 3169(0.95)

0.99 =

3041 𝑘𝑔

𝑑𝑎𝑦

𝐼 = 125 𝑘𝑔

𝑑𝑎𝑦

147

MILLING

ASSUMPTION

1.0% of the crystals are assumed to go waste during the milling process.

CALCULATIONS

Total mass balance on milling

𝑇𝑜𝑡𝑎𝑙 𝑚𝑎𝑠𝑠 𝑖𝑛 = 𝑡𝑜𝑡𝑎𝑙 𝑚𝑎𝑠𝑠 𝑜𝑢𝑡

𝐽 = 𝐾 + 𝐿

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

𝐴𝑚𝑜𝑢𝑛𝑡 𝑎𝑠 𝑤𝑎𝑠𝑡𝑒 (𝐾) = 1.0

100 ×

3041𝑘𝑔

𝑑𝑎𝑦

𝐴𝑚𝑜𝑢𝑛𝑡 𝑎𝑠 𝑤𝑎𝑠𝑡𝑒 (𝐾) ≅30.41 𝑘𝑔

𝑑𝑎𝑦

148

𝐹𝑟𝑜𝑚 𝑡ℎ𝑒 𝑡𝑜𝑡𝑎𝑙 𝑚𝑎𝑠𝑠 𝑏𝑎𝑙𝑎𝑛𝑐𝑒, 𝐿 =3041𝑘𝑔

𝑑𝑎𝑦 −

30.41 𝑘𝑔

𝑑𝑎𝑦

𝑇ℎ𝑒 𝑓𝑙𝑜𝑤 𝑟𝑎𝑡𝑒 𝑜𝑓 𝑓𝑖𝑛𝑒 𝑠𝑜𝑟𝑏𝑖𝑡𝑜𝑙 𝑝𝑜𝑤𝑑𝑒𝑟 (𝐿) ≈3000 𝑘𝑔

𝑑𝑎𝑦

PACKAGING

ASSUMPTION

No mass loss or gain during packaging.

CALCULATION

Output of the plant = 3000kg of sorbitol daily.

Quantity in 50kg bags = 3000 𝑘𝑔

50 𝑘𝑔

𝑄𝑢𝑎𝑛𝑡𝑖𝑡𝑦 𝑜𝑓 𝑠𝑜𝑟𝑏𝑖𝑡𝑜𝑙 = 60 𝑏𝑎𝑔𝑠

𝑑𝑎𝑦

𝑇ℎ𝑒𝑟𝑒𝑓𝑜𝑟𝑒 𝑓𝑜𝑟 𝑎 𝑓𝑒𝑒𝑑 𝑜𝑓5023𝑘𝑔

𝑑𝑎𝑦, 60 𝑏𝑎𝑔𝑠 𝑜𝑓 𝑎 50𝑘𝑔 𝑠𝑜𝑟𝑏𝑖𝑡𝑜𝑙 𝑐𝑟𝑦𝑠𝑡𝑎𝑙𝑠 𝑤𝑖𝑙𝑙 𝑏𝑒 𝑜𝑏𝑡𝑎𝑖𝑛𝑒𝑑.

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.

∆𝐸 = ��(∆𝐻 + ∆𝐾𝐸 + ∆𝑃𝐸) ± 𝑄 ± 𝑊

At steady state, ∆𝐸 = 0

There is no movement of equipment, velocity is zero hence ∆𝐾𝐸 = 0

∆𝑃𝐸 = 0

The above equation reduces to

𝑄 = ��∆𝐻 + 𝑊

��∆𝐻 = 𝐹𝑠𝐶𝑝𝑠∆𝑇

(Siebel 1892), proposed the following formula for estimating the specific heat values above and

below freezing;

𝐶𝑝 = 3.35𝑋𝑤 + 0.84 (𝐾𝐽/𝑘𝑔𝐾)………………………( 𝑎𝑏𝑜𝑣𝑒 𝑓𝑟𝑒𝑒𝑧𝑖𝑛𝑔)

𝐶𝑝 = 1.26𝑋𝑤 + 0.84 (𝐾𝐽

𝑘𝑔𝐾)………………………( 𝑏𝑒𝑙𝑜𝑤 𝑓𝑟𝑒𝑒𝑧𝑖𝑛𝑔)

Where 𝑋𝑤 = 𝑡ℎ𝑒 𝑚𝑎𝑠𝑠 𝑟𝑎𝑡𝑖𝑜 𝑜𝑓 𝑤𝑎𝑡𝑒𝑟 𝑖𝑛 𝑡ℎ𝑒 𝑝𝑟𝑜𝑑𝑢𝑐𝑡

150

Using equation (1)

The specific capacity of the starch slurry at 70% of moisture content; that is 𝑋𝑤 = 0.7

𝐶𝑝𝑠 = 3.35(0.7) + 0.84 (𝐾𝐽

𝐾𝑔𝐾)

𝐶𝑝𝑠 = 3.185 (𝐾𝐽

𝑘𝑔𝐾)

𝐶𝑝𝑠 𝑖𝑠 𝑡ℎ𝑒 𝑠𝑝𝑒𝑐𝑖𝑓𝑖𝑐 ℎ𝑒𝑎𝑡 𝑐𝑎𝑝𝑎𝑐𝑖𝑡𝑦 𝑜𝑓 𝑡ℎ𝑒 𝑠𝑡𝑎𝑟𝑐ℎ 𝑠𝑙𝑢𝑟𝑟𝑦

��∆𝐻 = 𝐹𝑠𝐶𝑝𝑠∆𝑇

��∆𝐻 = 6048.8 × 3.185 × (97 − 25)

= 1483437𝐾𝐽

𝑑𝑎𝑦(

1𝑑𝑎𝑦 𝑥 1ℎ𝑟

24 ℎ𝑟 × 3600𝑠)

= 17.2 𝑘𝑊

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 λs

𝑚𝑠 = 11.5

𝛌𝐬

151

𝑓𝑟𝑜𝑚 𝑠𝑡𝑒𝑎𝑚 𝑡𝑎𝑏𝑙𝑒, 𝑠𝑡𝑒𝑎𝑚 𝑎𝑡 120℃ ℎ𝑎𝑠 𝛌𝐬 = 𝟐𝟐𝟓𝟖

=17.2

2258= 7.6 × 10−3 kg/s

= (7.6 × 10−3) 𝑘𝑔

𝑠(3600𝑠)

1ℎ𝑟

𝑚𝑠= 27.4 𝑘𝑔/ℎ𝑟

Therefore18.4 kg/hr of steam is needed to supply this energy

SURFACE AREA REQUIRED

𝑄 = 𝑈𝐴𝑠∆𝑇

𝑄 = 17.2 𝑘𝑊

𝑈 = 𝑜𝑣𝑒𝑟𝑎𝑙𝑙 ℎ𝑒𝑎𝑡 𝑡𝑟𝑎𝑛𝑠𝑓𝑒𝑟 𝑐𝑜𝑒𝑓𝑓𝑖𝑐𝑖𝑒𝑛𝑡 = 300 𝑊/𝑚2℃

Mean liquid temperature 𝑇𝑚 =𝑇1+𝑇2

2=

97+25

2

= 61℃

152

Steam at 100 kpa has temperature, 𝑇𝑠 = 120℃

∆𝑇 = Mean temperature difference = 𝑇𝑠 − 𝑇𝑚

= 120 − 61

∆𝑇 = 59℃

𝐴𝑠 =𝑄

∆𝑇𝑈

=17.2 ×103

59 ×300

𝐴𝑠 = 0.97 𝑚2

The work done by the stirrer, W

𝐷

𝑇= 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(𝑁𝑝) = 1.370

Power(P) = Np × ρ × 𝑁3 × 𝐷5

Power(P) = 1.370 × 1200 × 53.203 × 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

𝑄 = ��∆𝐻 + 𝑊

𝑄 = 17.2 + 1.38

𝑄 = 18.6 𝑘𝑊

The total amount of energy needed during liquefaction is 18.6 𝑘𝑊

SACCHARIFICATION PROCESS

During this process, the partially hydrolysed starch at 97°𝐶 is reduced to 60°𝐶, thus, there is heat

energy loss.

𝐶𝑝 = 3.35𝑋𝑤 + 0.84 (𝐾𝐽/𝑘𝑔𝐾)………………………( 𝑎𝑏𝑜𝑣𝑒 𝑓𝑟𝑒𝑒𝑧𝑖𝑛𝑔)

𝐶𝑝𝑆𝑜1= 3.35 × 0.50 + 0.84 (𝐾𝐽/𝑘𝑔𝐾)

𝐶𝑝𝑆𝑜1=2.52 (𝐾𝐽/𝑘𝑔°𝐶)

The specific heat capacity of the partially hydrolysed starch, 𝐶𝑝𝑆𝑜1 is 2.68 (𝐾𝐽/𝑘𝑔°𝐶)

154

The flow rate of partially hydrolysed starch, 𝑆𝑜1 = 4851

��∆𝐻 = 𝑆𝑜1 × 𝐶𝑝𝑆𝑜1× ∆𝑇

= 4851 × 2.52 × (60 − 97)

= −452307𝐾𝐽

𝑑𝑎𝑦(

1𝑑𝑎𝑦 𝑥1ℎ𝑟

24ℎ𝑟 𝑥 3600𝑠)

= −5.2 𝑘𝑊

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 28°𝐶 and leaves at 38°𝐶

𝑄 = 𝑚𝑤𝐶𝑝𝑤∆𝑇

𝑚𝑤 =𝑄

𝐶𝑝𝑤∆𝑇

=5.2

4.187(38−28)

=0.12 kg/s

= 430 𝑘𝑔/ℎ𝑟

Therefore, 430 𝑘𝑔/ℎ𝑟 of water at 28°𝐶 is needed.

155

AREA OF STEAM JACKET REQUIRED

𝑄 = 𝑈𝐴𝑠∆𝑇

𝑄 = 5.2 𝑘𝑊

𝑈 = 𝑜𝑣𝑒𝑟𝑎𝑙𝑙 ℎ𝑒𝑎𝑡 𝑡𝑟𝑎𝑛𝑠𝑓𝑒𝑟 𝑐𝑜𝑒𝑓𝑓𝑖𝑐𝑖𝑒𝑛𝑡 = 300 𝑊/𝑚2℃

Mean liquid temperature 𝑇𝑚 =𝑇1+𝑇2

2=

97+60

2

= 79.5℃

Steam at 100 kpa has temperature, 𝑇𝑠 = 120℃

∆𝑇 = Mean temperature difference = 𝑇𝑠 − 𝑇𝑚

= 120 − 79.5

∆𝑇 = 41.5℃

𝐴𝑠 =𝑄

∆𝑇𝑈

=5.2 × 103

41.5 × 300

𝐴𝑠 = 0.4 𝑚2

The work done by the propeller, W

𝐷

𝑇= 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(𝑁𝑝) = 1.370

Power(P) = Np × ρ × 𝑁3 × 𝐷5

Power(P) = 1.370 × 1540 × 60.903 × 867.085

Power(P) = 1.08 kW

Assuming loading of 80% efficiency, motor power required

Motor Power = 1.08

= 1.35 kW

Hence

𝑄 = ��∆𝐻 + 𝑊

𝑄 = −5.2 + 1.35

𝑄 = −3.85 𝑘𝑊

There are -3.85𝑘𝑊 of energy lost during saccharification.

157

PLATE HEAT EXCHANGER WITH FOUR PLATES

𝑋𝑤 = 0.5

𝑝 = 3.35𝑋𝑤 + 0.84 (𝐾𝐽/𝑘𝑔𝐾)………………………( 𝑎𝑏𝑜𝑣𝑒 𝑓𝑟𝑒𝑒𝑧𝑖𝑛𝑔)

𝐶𝑝𝐺02= 3.35 × 0.5 + 0.84 (𝐾𝐽/𝑘𝑔𝐾)

𝐶𝑝𝐺02=2.52 (𝐾𝐽/𝑘𝑔°𝐶)

The specific heat capacity of the partially hydrolysed starch, 𝐶𝑝𝑆𝑜1 is 2.52 (𝐾𝐽/𝑘𝑔°𝐶)

𝑄 = 𝐺𝑜2 × 𝐶𝑝𝐺01× ∆𝑇

= 𝐺𝑜2 × 𝐶𝑝𝐺01× (𝑇2 − 𝑇1)

𝑄 = 5103.9 × 2.52 × (77 − 60)

= 218651𝐾𝐽

𝑑𝑎𝑦(

1𝑑𝑎𝑦 × 1ℎ𝑟

24 𝑑𝑎𝑦×3600𝑠)

= 2.5 𝑘𝑊

158

The water enters at a temperature of 𝑇𝑤1 = 110°𝐶 and leaves at 𝑇𝑤2 =95°𝐶

𝑄 = 𝑚𝑤𝐶𝑝𝑤∆𝑇

𝑚𝑤 =𝑄

𝐶𝑝𝑤(𝑇𝑤1−𝑇𝑤2 )

=2.5

4.187(110−95)

=0.04 kg/s

= 145 𝑘𝑔/ℎ𝑟

Therefore, 89.7 𝑘𝑔/ℎ𝑟 of water at °𝐶 is needed.

∆𝑇𝑚 =(𝑇𝑤2 −𝑇1)−(𝑇𝑤1 −𝑇2)

ln [(𝑇𝑤2 −𝑇1)

(𝑇𝑤1 −𝑇2)]

=(95−60)−(110−77)

ln [(95−601)

(110−77)]

= 34°𝐶

159

Calculating the Area of the heat exchanger

𝑄 = 𝑈𝐴𝑠∆𝑇𝑚

The Overall heat transfer coefficient 𝑈 = 4.23 𝑘𝑊/𝑚2𝐾

𝐴𝑠 =2.5

4.23×34

= 0.02 𝑚2

Therefore, the area of the heat exchanger is 0.02 𝑚2

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 ℃. 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.

∆𝐸 = ��(∆𝐻 + ∆𝐾𝐸 + ∆𝑃𝐸) ± 𝑄 ± 𝑊

At steady state, ∆𝐸 = 0

There is no movement of equipment, ∆𝐾𝐸 = 0

Since there are no elevations, ∆𝑃𝐸 = 0

Since there is no work done, 𝑊 = 0

The above equation reduces to

𝑄 = 𝑚∆𝐻

161

DATA

Temperature inlet, T1 = 70 ℃

Temperature outlet T2 = 100℃

Temperature of steam Ts = 110 ℃

𝑓𝑟𝑜𝑚 𝑄 = 𝑚∆𝐻

𝑄 = 5023 × 2.68 × (100 − 70)

𝑄 = 403849 𝑘𝐽

This means the quantity of heat needed by the system is 403849 𝒌𝑱

𝑞𝑆 = 403849 𝑘𝐽

24 × 3600 𝑠= 4.67 𝑘𝐽

The mean heat transfer rate, 𝒒𝑺 = 4.67 kW

∆𝑇𝑚 =(𝑇𝑠 − 𝑇1) − (𝑇𝑠 − 𝑇2)

ln [(𝑇𝑠 − 𝑇1)(𝑇𝑠 − 𝑇2)

]

∆𝑇𝑚 =(110 − 70) − (110 − 100)

ln [(110 − 70)(110 − 100)

]

∆𝑇𝑚 = 21.6 ℃

162

Using the equation 𝑄 = 𝐴𝑈∆𝑇𝑚

Taking the overall heat transfer coefficient U = 0.50 kW /m2K

𝐴 =𝑄

𝑈∆𝑇𝑚=

4.67

0.5 × 21.6

𝐴 = 𝟎. 𝟒𝟑 m2

Mass of steam

𝑚 = 𝑞

𝐶𝑝∆𝑇=

4.67

2.68 × (100 − 70)

𝑚 = 222.4 𝑘𝑔/ℎ

The mass of steam needed to raise the temperature from 70 ℃ to 100℃ using th shell and tube

heat exchanger is 𝟐𝟐𝟐. 𝟒 𝒌𝒈/𝒉

HYDROGENATION

The energy balance on the equipment is based on the law of conservation of energy.

∆𝐸 = ��(∆𝐻 + ∆𝐾𝐸 + ∆𝑃𝐸) ± 𝑄 ± 𝑊

At steady state, ∆𝐸 = 0

There is no movement of equipment, ∆𝐾𝐸 = 0

Since there are no elevations, ∆𝑃𝐸 = 0

The above equation reduces to

𝑄 = 𝑚∆𝐻 + 𝑊

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 𝑄 = 𝑄𝐺+𝑄𝑆

DATA FOR CALCULATING THE HEAT ADDED TO THE SYSTEM (QS)

Mass flow rate of sorbitol, m =5474kg/day

Inlet temperature, T1 = 100 ℃

Outlet temperature T2 = 130 ℃

Specific heat capacity of sorbitol using Seibel, 𝐶𝑝 = 0.875

Enthalpy of inlet steam at 100 ℃ , H1 =2676kJ/kg

Enthalpy of outlet steam at 130℃, H2 = 2787 kJ/kg

𝑄𝑆 = 𝑚∆𝐻 + 𝑊

𝑚∆𝐻 = 5474 × (2787 − 2676) + 𝑊

𝑸𝑺 = 𝟔𝟎𝟕𝟔𝟏𝟒 + 𝑾

164

THE WORK DONE BY THE STIRRER, W

The ratios of the diameter of the stirrer (𝐷𝑆) to the diameter of the tank (𝐷𝑇) range from 0.3 to

12. Taking a ratio of 0.5. But the tank has a diameter 𝐷𝑇 of 1.8m

𝐷𝑆

𝐷𝑇= 0.5

𝐷𝑆 = 0.5 × 1.8 = 0.90 𝑚

DATA

Density of sorbitol, = 1490 kg/m3

Agitator Speed (N) = 8.4𝑚𝑠−1 = 180 rpm

Power Number(𝑁𝑝) = 1.370

Using the equation, Power(P) = Np × ρ × 𝑁3 × 𝐷5

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,𝐖 = 0.72 × 24 × 3600 = 62208 kJ

Substituting, W= 62208 kJ

𝑄𝑆 = 607614 + 𝑊

𝑄𝑆 = 607614 + 62208

𝑸𝑺 = 𝟔𝟔𝟗𝟖𝟐𝟐 kJ.

165

This means the quantity of heat needed by the system is 𝟔𝟔𝟗𝟖𝟐𝟐 𝐤J.

𝑞𝑆 = 669822 𝑘𝐽

24 × 3600 𝑠

𝒒𝑺 = 𝟕. 𝟕𝟓 𝒌𝑾

The mean heat transfer rate, 𝒒𝑺 = 𝟕. 𝟕𝟓 kW

The quantity of steam needed to provide this energy 𝑚𝑠

𝑚𝑠 =��∆𝐻

λs 𝑤ℎ𝑒𝑟𝑒 λs at 130 ℃ = 2787kJ/kg

𝑚𝑠 =7.75

2787 = 9.08𝑘𝑔/ℎ

Therefore, the amount of steam required to supply the energy is = 9.08 𝒌𝒈/𝒉

HEAT GENERATED BY THE SYSTEM (QG)

𝐶6𝐻12𝑂6 + 𝐻2 − − − 𝐶6𝐻14𝑂6

𝐸𝑛𝑡ℎ𝑎𝑙𝑝𝑦 𝑜𝑓𝑓𝑜𝑟𝑚𝑎𝑡𝑖𝑜𝑛 𝑜𝑓 𝑔𝑙𝑢𝑐𝑜𝑠𝑒, ∆𝐻𝐶 (𝐶6𝐻12𝑂6) = 2805 𝑘𝐽𝑚𝑜𝑙−1

𝐸𝑛𝑡ℎ𝑎𝑙𝑝𝑦 𝑜𝑓 𝑓𝑜𝑟𝑚𝑎𝑡𝑖𝑜𝑛 𝑜𝑓 ℎ𝑦𝑑𝑟𝑜𝑔𝑒𝑛, ∆𝐻𝐶 (𝐻2) = −285.5 𝑘𝐽𝑚𝑜𝑙−1

𝐸𝑛𝑡ℎ𝑎𝑙𝑝𝑦 𝑜𝑓 𝑓𝑜𝑟𝑚𝑎𝑡𝑖𝑜𝑛 𝑜𝑓 𝑠𝑜𝑟𝑏𝑖𝑡𝑜𝑙, ∆𝐻𝑓 (𝐶6𝐻14𝑂6) = −3009 𝑘𝐽𝑚𝑜𝑙−1

𝑓𝑟𝑜𝑚 𝑯𝒆𝒔𝒔 𝒍𝒂𝒘,

∆𝐻𝑟𝑥𝑛 = ∑∆𝐻𝑝𝑟𝑜𝑑𝑢𝑐𝑡 − ∑∆𝐻𝑟𝑒𝑎𝑐𝑡𝑎𝑛𝑡

Where ∑∆𝐻𝑝𝑟𝑜𝑑𝑢𝑐𝑡 = 𝑒𝑛𝑡ℎ𝑎𝑙𝑝𝑦 𝑜𝑓 𝑠𝑜𝑟𝑏𝑖𝑡𝑜𝑙 𝑓𝑜𝑟𝑚𝑒𝑑

166

∑∆𝐻𝑟𝑒𝑎𝑐𝑡𝑎𝑛𝑡𝑠 = 𝑒𝑛𝑡ℎ𝑎𝑙𝑝𝑦 𝑜𝑓 𝑔𝑙𝑢𝑐𝑜𝑠𝑒 + 𝑒𝑛𝑡ℎ𝑎𝑙𝑝𝑦 𝑜𝑓 ℎ𝑦𝑑𝑟𝑜𝑔𝑒𝑛 𝑟𝑒𝑎𝑐𝑡𝑒𝑑

∆𝐻𝑟𝑥𝑛 = −3009 − (2805 − 285.5)

∆𝐻𝑟𝑥𝑛 = −489.9 𝑘𝐽𝑚𝑜𝑙−1

𝐵𝑢𝑡 𝑡ℎ𝑒 𝑚𝑜𝑙𝑒𝑠, 𝑛 𝑜𝑓 𝑠𝑜𝑟𝑏𝑖𝑡𝑜𝑙 = 30.05 𝑚𝑜𝑙 𝑎𝑛𝑑 𝑡ℎ𝑒 𝑚𝑜𝑙𝑎𝑟 𝑚𝑎𝑠𝑠 ,𝑀 = 182.17

𝐶𝑜𝑛𝑣𝑒𝑟𝑡𝑖𝑛𝑔 𝑡ℎ𝑒 ∆𝐻𝑟𝑥𝑛𝑓𝑟𝑜𝑚 𝑘𝐽𝑚𝑜𝑙−1 𝑡𝑜 𝑘𝐽𝑘𝑔−1 = −2.73 𝑘𝐽/𝑘𝑔

𝑄𝐺 = 𝑚 ∆𝐻

𝑄𝐺 = 5474(−2.73)

𝑸𝑮 = −𝟏𝟒𝟗𝟒𝟏. 𝟗𝟓 𝒌𝑱

The quantity of energy generated by the system is = −14941.95 𝑘𝐽.

𝑞𝐺 =−14941.95

24 × 3600= −𝟎. 𝟏𝟕𝟑𝒌𝑾

Substituting the values of 𝑸𝑺 𝒂𝒏𝒅 𝑸𝑮

𝑄 = 𝑄𝐺+𝑄𝑆

𝑸 = 669822 − 14941.95 = 𝟔𝟓𝟒𝟖𝟖𝟎 𝒌𝑱

The total amount of energy in the hydrogenation reactor is 𝟔𝟓𝟒𝟖𝟖𝟎 𝒌𝑱

𝒒 =𝟔𝟓𝟒𝟖𝟖𝟎 𝑘𝐽

24×3600= 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 25℃ and leave at 53℃

𝑄 = 𝑚𝐶𝑝∆𝑇

Where the heat capacity of water 𝐶𝑝 = 4.18𝑘𝐽𝑘𝑔−1

𝑚 =−𝟏𝟒𝟗𝟒𝟏.𝟗𝟓

4.18×(25−53)= 4.79 𝑘𝑔/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

𝑏𝑢𝑡 𝑎𝑡 1 𝑎𝑡𝑚 = 760𝑚𝑚𝐻𝑔

𝑡ℎ𝑒𝑟𝑒𝑓𝑜𝑟𝑒 75𝑚𝑚𝐻𝑔 =75

760× 1𝑎𝑡𝑚

= 0.0987 𝑎𝑡𝑚

𝑐𝑜𝑛𝑣𝑒𝑟𝑡𝑖𝑛𝑔 𝑡ℎ𝑖𝑠 𝑡𝑜 𝑘𝑃𝑎 𝑢𝑠𝑖𝑛𝑔 1𝑎𝑡𝑚 = 101.325, 𝑃𝑟𝑒𝑠𝑠𝑢𝑟𝑒 𝑖𝑠 10.00 𝑘𝑃𝑎

𝐹𝑟𝑜𝑚 𝑠𝑡𝑒𝑎𝑚 𝑡𝑎𝑏𝑙𝑒, 𝑎 𝑝𝑟𝑒𝑠𝑠𝑢𝑟𝑒 𝑜𝑓 10.00 𝑘𝑃𝑎 𝑐𝑜𝑟𝑟𝑒𝑠𝑝𝑜𝑛𝑑𝑠 𝑡𝑜 𝑎 𝑡𝑒𝑚𝑝𝑒𝑟𝑎𝑡𝑢𝑟𝑒 𝑜𝑓 46℃

168

DATA

Initial temperature of sorbitol, T1= 30 ℃

Final temperature of sorbitol, T2 = 46 ℃

Cp of sorbitol using the Siebel equation =0.875 kJ/KgK

𝑈𝑠𝑖𝑛𝑔 𝑄 = 𝑚𝐶𝑃∆𝑇

𝑄 = 4563.36𝑘𝑔 × 0.875 × (46 − 30)

𝑸 = 𝟔𝟑𝟖𝟖𝟕. 𝟎𝟒𝒌𝑱

The quantity of heat energy needed is 63887.04𝑘𝐽

𝒒 =63887.04

24 × 3600= 𝟎. 𝟕𝟑𝟗 𝒌𝑾

Therefore, the heat transfer rate is 0.74kW

Quantity of steam needed.

𝑚𝑆 =𝑞

ℎ𝑒

Where he is the latent heat of evaporation at 45C = 2393 kJ/kg

𝑀𝑠 =0.74

239= 3.089 × 104

𝑀𝑠 = 1.11𝑘𝑔/ℎ

𝑡ℎ𝑒 𝑚𝑎𝑠𝑠 𝑜𝑓 𝑠𝑡𝑒𝑎𝑚 𝑟𝑒𝑞𝑢𝑖𝑟𝑒𝑑 𝑡𝑜 𝑝𝑟𝑜𝑣𝑖𝑑𝑒 𝑡ℎ𝑒 𝑒𝑛𝑒𝑟𝑔𝑦 𝑖𝑠 𝟏. 𝟏𝟏 𝒌𝒈/𝒉

169

ENERGY BALANCE ON CRYSTALLISATION

The energy balance on the various stages and equipment is based on the law of conservation of

energy.

∆𝐸 = ��(∆𝐻 + ∆𝐾𝐸 + ∆𝑃𝐸) ± 𝑄 ± 𝑊

At steady state, ∆𝐸 = 0

There is no movement of equipment, ∆𝐾𝐸 = 0

Since there are no elevations, ∆𝑃𝐸 = 0

Work done W =0

The above equation reduces to

𝑄 = 𝑚∆𝐻

DATA

Initial temperature of sorbitol, T1 = 40 ℃

Final temperature of sorbitol, T2 = 95℃

Mass flow rate of sorbitol through the crystallizer, m = 4301 kg/day

𝑄 = 4301 × 0.875 × (95 − 40)

𝑄 = 206986.6

𝑞 =206986.6

24 × 3600= 2.40𝑘𝑊

170

MASS OF STEAM REQUIRED TO PRODUCE THIS ENERGY

𝑚𝑆 =𝑞

ℎ𝑒 𝑤ℎ𝑒𝑟𝑒 ℎ𝑒 𝑡ℎ𝑒 𝑙𝑎𝑡𝑒𝑛𝑡 ℎ𝑒𝑎𝑡 𝑜𝑓 𝑐𝑟𝑦𝑠𝑡𝑎𝑙𝑙𝑖𝑠𝑎𝑡𝑖𝑜𝑛 𝑎𝑡 90 ℃ = 2270

𝑘𝐽

𝑘𝑔

𝑚𝑆 =2.40

2270= 0.10

𝑘𝑔

ℎ

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

𝑄 = 𝑄𝑤𝑒 + 𝑄𝑠ℎ + 𝑄𝑎ℎ

𝑄𝑤𝑒 = 𝐹(𝑋𝑜 − 𝑋)[∆𝐻𝑜 − (𝐶𝑃𝐿 − 𝐶𝑃𝑉)𝑇]

𝑄𝑠ℎ = 𝐹(𝐶𝑃𝑆 + 𝑋𝑜𝐶𝑃𝐿)(𝑇 − 𝑇𝑜)

𝑄𝑎ℎ = 𝐹𝑎[𝐶𝑃𝐴 + 𝑌𝑜𝐶𝑃𝑉](𝑇 − 𝑇𝑜), where

𝐹𝑎 = 𝐹(𝑋𝑜 − 𝑋)

𝑌 − 𝑌𝑜

171

Where:

F= flow rate of sorbitol crystals (kg/h) =3169

𝑑𝑎𝑦= 132.04 𝑘𝑔/ℎ

𝐹𝑎 (𝑘𝑔

ℎ)= flow rate of fresh air entering dryer= 12 kg/h

𝑋𝑜=𝑋𝑤 = % moisture content of sorbitol crystals = 0.05

X= % moisture content of dried sorbitol crystals= 0.01

𝑌𝑜= ambient humidity= 0.01

Y= drying air humidity= 0.45 kg/kg dry basis

∆𝐻𝑜= Latent heat at 0oC= 2500 kJ/kg

𝐶𝑃𝐿= specific heat capacity of water= 4.2 kJ/kg k

𝐶𝑃𝑉= specific heat capacity of vapor= 1.90 kJ/kg k

𝐶𝑃𝑆= specific heat capacity of sorbitol=0.875𝑘𝐽/𝑘𝑔 𝑘

T= drying air temperature= 65oC

To= ambient temperature= 25oC

172

Calculations

F= 3169 𝑘𝑔

𝑑𝑎𝑦×

1 𝑑𝑎𝑦

24 ℎ𝑟𝑠= 132.04𝑘𝑔/ℎ𝑟

𝐶𝑃𝑆 𝑢𝑠𝑖𝑛𝑔 𝑡ℎ𝑒 𝑓𝑜𝑟𝑚𝑢𝑙𝑎𝑒 𝑝𝑟𝑜𝑝𝑜𝑠𝑒𝑑 𝑏𝑦 Siebel (1892) 𝑓𝑜𝑟 estimating the specific heat values

above and below freezing;

𝐶𝑝 = 3.35𝑋𝑤 + 0.84 (𝐾𝐽/𝑘𝑔𝐾)………………………( 𝑎𝑏𝑜𝑣𝑒 𝑓𝑟𝑒𝑒𝑧𝑖𝑛𝑔)

𝐶𝑝 = 1.26𝑋𝑤 + 0.84 (𝐾𝐽

𝑘𝑔𝐾)………………………( 𝑏𝑒𝑙𝑜𝑤 𝑓𝑟𝑒𝑒𝑧𝑖𝑛𝑔)

Where 𝑋𝑤 = 𝑡ℎ𝑒 𝑚𝑎𝑠𝑠 𝑟𝑎𝑡𝑖𝑜 𝑜𝑓 𝑤𝑎𝑡𝑒𝑟 𝑖𝑛 𝑡ℎ𝑒 𝑝𝑟𝑜𝑑𝑢𝑐𝑡

𝐶𝑝 = 3.5(0.01) + 0.84 = 0.875 𝑘𝐽/𝑘𝑔 𝑘

𝐹𝑎 =132.04(0.05 − 0.01)

0.45 − 0.01= 12.00𝑘𝑔/ℎ

Thermal Requirements

𝑄𝑤𝑒 = 132.04(0.05 − 0.0)[2500 − (4.2 − 1.90)(65)]

𝑄𝑤𝑒 ≈12414.4𝑘𝐽

ℎ×

1ℎ

3600𝑠≈ 3.45𝑘𝑊

𝑄𝑠ℎ = 132.04(0.875 + 0.05(4.2))(65 − 25)

𝑄𝑠ℎ =5730.547

ℎ×

1ℎ

3600𝑠≈ 1.59𝑘𝑊

𝑄𝑎ℎ = 12[1.0 + (0.01)(1.90)](65 − 25)

173

𝑄𝑎ℎ ≈407.6

ℎ×

1ℎ

3600𝑠≈ 0.135 𝑘𝑊

Total Thermal Requirement for drying

𝑄 = 3.45 + 1.59 + 0.135 = 5.18𝑘𝑊

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/𝑚3

Velocity of the laminar flow 𝑉 = 2.0 m/s

Viscosity of starch at 25 ℃, 𝜇 = 3.5 × 10−3 Pas

Specific gravity, 𝑆𝑔=

PRESSURE DROP

∆𝑃 =𝑓𝑙𝑉2𝑆𝑔

2𝑑

𝐵𝑢𝑡 𝑓 =64

𝑅𝑒

𝐵𝑢𝑡 𝑅𝑒 =𝑣𝑑

=

893 × 2 × 82.9 × 10−3

3.5 × 10−3= 42302.7

𝑓 =64

42302.7= 1.5 × 10−3

175

𝑇ℎ𝑒𝑟𝑒𝑓𝑜𝑟𝑒 ∆𝑃 =1.5 × 10−3 × 2 × 4 × 893 × 0.893

2 × 82.9 × 10−3= 57.72 𝑃𝑎

MIXING TANK TO EXTRACTOR

Diameter of pipe, d = 82.9 mm

Length of pipe, L = 5m

Density of starch pulp, = 826 Kg/𝑚3

Velocity of the laminar flow 𝑉 = 2.0 m/s

Viscosity of starch at 25 ℃, 𝜇 = 3.5 × 10−3 Pa.s

Specific gravity, 𝑆𝑔= 0.826

PRESSURE DROP

∆𝑃 =𝑓𝑙𝑉2𝑆𝑔

2𝑑

𝐵𝑢𝑡 𝑓 =64

𝑅𝑒

𝐵𝑢𝑡 𝑅𝑒 =𝑉𝑑

=

826 × 2 × 82.9 × 10−3

3.5 × 10−3= 39128.8

176

𝑓 =64

39128.8= 1.64 × 10−3

𝑇ℎ𝑒𝑟𝑒𝑓𝑜𝑟𝑒 ∆𝑃 =1.64 × 10−3 × 5 × 4 × 826 × 0.826

2 × 82.9 × 10−3= 134.97 𝑃𝑎

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/𝑚3

Velocity of the laminar flow 𝑉 = 2.0 m/s

Viscosity of starch at 25 ℃, 𝜇 = 3.5 × 10−3 Pa.s

Specific gravity, 𝑆𝑔= 0.826

PRESSURE DROP

∆𝑃 =𝑓𝑙𝑉2𝑆𝑔

2𝑑

𝐵𝑢𝑡 𝑓 =64

𝑅𝑒

177

𝐵𝑢𝑡 𝑅𝑒 =𝑉𝑑

=

826 × 2 × 66.7 × 10−3

3.5 × 10−3= 62941.2

𝑓 =64

62941.2= 1.02 × 10−3

𝑇ℎ𝑒𝑟𝑒𝑓𝑜𝑟𝑒 ∆𝑃 =1.02 × 10−3 × 4 × 4 × 826 × 0.826

2 × 66.7 × 10−3= 83.47 𝑃𝑎

FROM SEPARATOR TO DEWATERING EQUIPMENT

Diameter of pipe, d = 82.9 mm

Length of pipe, L = 6m

Density of starch pulp, = 893 Kg/𝑚3

Velocity of the laminar flow 𝑉 = 2.0 m/s

Viscosity of starch at 25 ℃, 𝜇 = 3.5 × 10−3 Pa.s

Specific gravity, 𝑆𝑔= 0.893

PRESSURE DROP

∆𝑃 =𝑓𝑙𝑉2𝑆𝑔

2𝑑

178

𝐵𝑢𝑡 𝑓 =64

𝑅𝑒

𝐵𝑢𝑡 𝑅𝑒 =𝑉𝑑

=

893 × 2 × 82.9 × 10−3

3.5 × 10−3= 42302.7

𝑓 =64

42302.7= 1.51 × 10−3

𝑇ℎ𝑒𝑟𝑒𝑓𝑜𝑟𝑒 ∆𝑃 =1.51 × 10−3 × 6 × 4 × 893 × 0.893

2 × 82.9 × 10−3= 174.30 𝑃𝑎

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/𝑚3

Velocity of the laminar flow 𝑉 = 2.0 m/s

Viscosity of starch at 25 ℃, 𝜇 = 3.5 × 10−3 Pas

Specific gravity, 𝑆𝑔= 0.81

PRESSURE DROP

∆𝑃 =𝑓𝑙𝑉2𝑆𝑔

2𝑑

𝐵𝑢𝑡 𝑓 =64

𝑅𝑒

𝐵𝑢𝑡 𝑅𝑒 =𝑣𝑑

=

810 × 2 × 84.7 × 10−3

3.5 × 10−3= 39204

𝑓 =64

39204= 1.6 × 10−3

𝑇ℎ𝑒𝑟𝑒𝑓𝑜𝑟𝑒 ∆𝑃 =1.6 × 10−3 × 8 × 4 × 810 × 0.81

2 × 84.7 × 10−3= 198.30 𝑃𝑎

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/𝑚3

Velocity of the laminar flow 𝑉 = 2.0 m/s

Viscosity of glucose at 25 ℃, 𝜇 = 6.8 × 10−3 Pa.s

Specific gravity, 𝑆𝑔= 1.54

PRESSURE DROP

∆𝑃 =𝑓𝑙𝑉2𝑆𝑔

2𝑑

𝐵𝑢𝑡 𝑓 =64

𝑅𝑒

𝐵𝑢𝑡 𝑅𝑒 =𝑣𝑑

=

1540 × 2 × 45 × 10−3

6.8 × 10−3= 20382

𝑓 =64

20382= 3.1 × 10−3

𝑇ℎ𝑒𝑟𝑒𝑓𝑜𝑟𝑒 ∆𝑃 =3.1 × 10−3 × 12 × 4 × 1540 × 1.54

2 × 45 × 10−3= 3921 𝑃𝑎

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/𝑚3

Velocity of the laminar flow 𝑉 = 2.0 m/s

Viscosity of glucose at 25 ℃, 𝜇 = 6.8 × 10−3 Pas

Specific gravity, 𝑆𝑔= 1.54

PRESSURE DROP

∆𝑃 =𝑓𝑙𝑉2𝑆𝑔

2𝑑

𝐵𝑢𝑡 𝑓 =64

𝑅𝑒

𝐵𝑢𝑡 𝑅𝑒 =𝑣𝑑

=

1540 × 2 × 54.7 × 10−3

6.8 × 10−3= 24776

𝑓 =64

24776= 2.6 × 10−3

𝑇ℎ𝑒𝑟𝑒𝑓𝑜𝑟𝑒 ∆𝑃 =2.6 × 10−3 × 10 × 4 × 1540 × 1.54

2 × 54.7 × 10−3= 36272 𝑃𝑎

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/𝑚3

Velocity of the laminar flow 𝑉 = 2.0 m/s

Viscosity of glucose at 25 ℃, 𝜇 = 6.8 × 10−3 Pas

Specific gravity, 𝑆𝑔= 1.54

PRESSURE DROP

∆𝑃 =𝑓𝑙𝑉2𝑆𝑔

2𝑑

𝐵𝑢𝑡 𝑓 =64

𝑅𝑒

𝐵𝑢𝑡 𝑅𝑒 =𝑣𝑑

=

1540 × 2 × 30.1 × 10−3

6.8 × 10−3= 13634

𝑓 =64

13634= 4.7 × 10−3

𝑇ℎ𝑒𝑟𝑒𝑓𝑜𝑟𝑒 ∆𝑃 =4.7 × 10−3 × 15 × 4 × 1540 × 1.54

2 × 30.1 × 10−3= 11096 𝑃𝑎

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/𝑚3

Velocity of the laminar flow 𝑉 = 2.0 m/s

Viscosity of glucose at 25 ℃, 𝜇 = 6.8 × 10−3 Pas

Specific gravity, 𝑆𝑔= 1.54

PRESSURE DROP

∆𝑃 =𝑓𝑙𝑉2𝑆𝑔

2𝑑

𝐵𝑢𝑡 𝑓 =64

𝑅𝑒

𝐵𝑢𝑡 𝑅𝑒 =𝑣𝑑

=

1540 × 2 × 54.7 × 10−3

6.8 × 10−3= 24776

𝑓 =64

24776= 2.6 × 10−3

𝑇ℎ𝑒𝑟𝑒𝑓𝑜𝑟𝑒 ∆𝑃 =2.6 × 10−3 × 15 × 4 × 1540 × 1.54

2 × 54.7 × 10−3= 3381 𝑃𝑎

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/𝑚3

Velocity of the laminar flow 𝑉 = 2.0 m/s

Viscosity of glucose at 25 ℃, 𝜇 = 6.8 × 10−3 Pa.s

Specific gravity, 𝑆𝑔= 1.54

PRESSURE DROP

∆𝑃 =𝑓𝑙𝑉2𝑆𝑔

2𝑑

𝐵𝑢𝑡 𝑓 =64

𝑅𝑒

𝐵𝑢𝑡 𝑅𝑒 =𝑣𝑑

=

1540 × 2 × 45 × 10−3

6.8 × 10−3= 20382

𝑓 =64

20382= 3.1 × 10−3

𝑇ℎ𝑒𝑟𝑒𝑓𝑜𝑟𝑒 ∆𝑃 =3.1 × 10−3 × 8 × 4 × 1540 × 1.54

2 × 45 × 10−3= 2614 𝑃𝑎

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

∆𝑃 =𝑓𝑙𝑉2𝑆𝑔

2𝑑

𝐵𝑢𝑡 𝑓 =64

𝑅𝑒

𝐵𝑢𝑡 𝑅𝑒 =𝑉𝑑

= 9060

𝑓 =64

9060= 7.06 × 10−3

186

𝑇ℎ𝑒𝑟𝑒𝑓𝑜𝑟𝑒 ∆𝑃 =7.06 × 10−3 × 3 × 4 × 1540 × 1.54

2 × 108.3 × 10−3= 0.93𝑃𝑎

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

∆𝑃 =𝑓𝑙𝑉2𝑆𝑔

2𝑑

𝐵𝑢𝑡 𝑓 =64

𝑅𝑒

𝐵𝑢𝑡 𝑅𝑒 =𝑣𝑑

= 867

𝑓 =64

867= 0.0738

𝑇ℎ𝑒𝑟𝑒𝑓𝑜𝑟𝑒 ∆𝑃 =0.0738 × 3 × 4 × 1490 × 1.49

2 × 135.7 × 10−3= 7.2 𝑘𝑃𝑎

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

∆𝑃 =𝑓𝑙𝑉2𝑆𝑔

2𝑑

𝐵𝑢𝑡 𝑓 =64

𝑅𝑒

𝐵𝑢𝑡 𝑅𝑒 =𝑣𝑑

= 867

𝑓 =64

867= 0.0738

𝑇ℎ𝑒𝑟𝑒𝑓𝑜𝑟𝑒 ∆𝑃 =0.0738 × 3 × 4 × 1490 × 1.49

2 × 134.5 × 10−3= 7.3 𝑃𝑎

189

APPENDIX C

DETAILED EQUIPMENT DESIGN CALCULATION

TOP SUSPENDED MOTOR CYLINDRICAL-SCREEN CENTRIFUGE

CALCULATIONS

CENTRIFUGAL FORCE

𝐶𝑒𝑛𝑡𝑟𝑖𝑓𝑢𝑔𝑎𝑙 𝑓𝑜𝑟𝑐𝑒, 𝐹𝑐 = 𝑚𝑟𝜔2

Where

𝐹𝑐= centrifugal force generated

𝑚 = mass of rotating particle (Kg)

𝜔 = angular speed (rad/min)

𝑟 = radius of rotation (m)

𝐹𝑐 = 𝑚𝑟𝜔2 =𝑚𝑢𝑡

2

𝑟

𝑢𝑡 =2𝜋𝑟𝑁

60

Hence

𝐹𝑐 = 𝑚

𝑟(2𝜋𝑟𝑁

60)2

190

But

𝑚 = total mass of basket content (magma) = 1000 kg

𝑁 = speed of rotation (rev/min) = 1200 rev/min

𝑟 = radius of rotation (m) = 0.6m

𝐹𝑐 = 𝑚

𝑟(2𝜋𝑟𝑁

60)2

= 1000

0.6(2𝜋 × 0.6 × 1200

60)

2

= 9474.8 𝐾𝑁

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.

𝑲𝑷 =𝝎𝟐

𝒈𝒓

𝐾𝑃 = separation factor

𝜔 = angular speed (rad/sec)

𝑔 = acceleration due to gravity (m/𝑠2) = 9.81 m/𝑠2

𝑟 = radius of rotation = 0.6m

𝑲𝑷 =𝝎𝟐

𝒈𝒓

191

But 𝜔 =2𝜋𝑁

60=

2𝜋×1200

60= 40𝜋

𝑲𝑷 =(40𝜋)2

9.81×0.6= 2682.9 ≅ 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 𝑲𝑷.

𝑪𝑭 = 𝑨 × 𝑲𝑷

But 𝑨 =𝝅(𝑫−𝒉)

𝑯

Where

𝑫 = diameter of basket = 1.2m

𝒉 = thickness of fluid on basket surface = 0.14m

𝑯 = length of cylindrical surface in contact with fluid = 1.0m

(𝑫 − 𝒉) = average diameter of rotation =𝑫+(𝑫−𝟐𝒉)

𝟐

Hence

𝑨 =𝝅(𝑫−𝒉)

𝑯=

𝝅×(𝟏.𝟐−𝟎.𝟏𝟒)

𝟏= 𝟑. 𝟑𝟑𝒎𝟐

192

𝑪𝑭 = 𝟑. 𝟑𝟑 × 2683 = 8934.39

= 8934.39 𝑚2

CAPACITY OF CENTRIFUGE

𝑄𝑡 = 𝑉𝑠 × 𝐶𝐹

𝑉𝑠 =1

18

𝑑𝑝𝑔(𝜌𝑝−𝜌𝑓)

𝜇

𝑅𝑒 = Reynolds number

𝜇 = viscosity of fluid = 199 Pa.s

𝑑𝑝 = diameter of particle = 0.002mm

𝜌𝑝 =149 Kg/𝑚3

𝑔 = 9.81 m/𝑠2

𝜌𝑓 = density of fluid = 1586.2 Kg/𝑚3

𝑉𝑠 =1

18

0.002 × 103 × 9.81 × (1586.2 − 149)

199= 7.9𝑚/𝑠

𝑄𝑡 = 𝑉𝑠 × 𝐶𝐹

= 7.9 × 8934.39 = 70581.7 𝑚3/𝑠

Capacity of centrifuge is 70581.7 𝑚3/𝑠

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;

𝜎𝑏 =𝑚𝑇𝜔2𝑟

𝜋𝐷𝐻=

𝑚𝑇𝜔2

𝜋𝐻

Where

𝜎𝑏 = stress on the walls of internal basket (N/𝑚2)

𝑚𝑇 = total mass of basket and its content (Kg) = 2110 Kg

𝐻 = height of basket (m) = 1.0m

𝐷 = Diameter of basket (m) = 1.2 m

𝜔 = 40𝜋

𝜎𝑏 =2110 × (40𝜋)2

𝜋 × 1= 10606 𝐾𝑁

𝜎𝑏 = 10606 𝐾𝑁

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:

𝑡𝑏 =𝜎𝑏𝐷

2𝜎𝑝

Where

𝑡𝑏= allowable thickness of basket material

𝜎𝑏 =10606 × 103

𝜎𝑝= permissible stress of the material of the basket = 115 × 106 𝑁/𝑚2

𝐷 = Diameter of basket (m) = 1.2m

𝑡𝑏 =𝜎𝑏𝐷

2𝜎𝑝=

10606 × 103 × 1.2

2 × 115 × 106= 0.06𝑚

𝑡𝑏 = 0.06𝑚

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.

𝑝𝑜𝑤𝑒𝑟 𝑟𝑒𝑞𝑢𝑖𝑟𝑒𝑚𝑒𝑛𝑡 = 𝐹𝑇 × 𝑉

Where

𝐹𝑇 = total force of basket, shaft and basket content (N)

𝐹𝑇 = 𝑚𝑇𝜔2𝑟

Assuming 𝑚𝑇 = 1000 + 1110 = 2110 𝐾𝑔

𝐹𝑇 = 2110 × (40𝜋)2 × 0.6 = 19991.9 𝐾𝑁

𝑉 = velocity of basket at full speed (m/sec) = 𝜔𝑟

𝑉 = 40𝜋 × 0.6 = 24𝜋 = 75.4𝑚

𝑠𝑒𝑐

𝑝𝑜𝑤𝑒𝑟 𝑟𝑒𝑞𝑢𝑖𝑟𝑒𝑚𝑒𝑛𝑡 = 75.4 × 19991.9 = 1507389.26 𝐾𝑊

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

𝑀𝑡 = twisting movement (Nm)

W = power transmitted (watts)

N = speed of rotation of the shaft (rev/sec)

𝑀𝑡 =60 × 1507389.26 × 103

2𝜋 ×120060

= 719725.4 𝐾𝑁𝑚

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).

𝑑 =(16𝑀𝑡𝐷)0.33

𝜋𝜎𝑝

197

𝑑 = diameter of shaft

𝑀𝑡 = twisting moment on shaft due to rotation of the internal basket fastened on shaft.

𝐷 = diameter of basket

𝜎𝑝 = permissible shear stress of stainless steel = 115 × 106 𝑁/𝑚2

𝑑 =(16×719725.4×103×1.2)0.33

𝜋×115×106= 6.1 × 10−6m

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

𝐺02 = 5104𝑘𝑔

𝑑𝑎𝑦

The volumetric flow rate 𝑄 =𝑚𝑎𝑠𝑠 𝑓𝑙𝑜𝑤 𝑟𝑎𝑡𝑒

𝑑𝑒𝑛𝑠𝑖𝑡𝑦

=5104

1540

= 3.3 𝑚3

𝑑𝑎𝑦

= 3314 𝐿/𝑑𝑎𝑦

The concentration of the colorants, ashes, and other impurities in the 3314 𝐿/𝑑𝑎𝑦 is 7.6 𝑚𝑔/𝐿 .

Assuming 90% removal, then the concentration remained in the flow stream is 0.76 𝑚 𝑔/𝐿

Liquefaction time is 90 minutes, and then the following parameters are chosen from above table

The Lagmuir isotherm parameter for NORIT; 𝑞𝑜, K and R2 values

Liquefaction

Time

𝑞𝑜 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

𝑞𝑜 = 344.83 𝑚𝑔/𝑔

K=26.93 L/mg

𝐶𝑒 = 0.76𝑚𝑔

𝐿

𝐶𝑒

𝑞𝑒=

1

𝑞𝑜𝐾1+

𝐶𝑒

𝑞𝑜

𝑞𝑒 = 𝑖𝑠 𝑡ℎ𝑒 𝑚𝑎𝑠𝑠 𝑜𝑓 𝑎𝑑𝑠𝑜𝑟𝑏𝑎𝑡𝑒 𝑎𝑑𝑠𝑜𝑟𝑏𝑒𝑑 𝑝𝑒𝑟 𝑢𝑛𝑖𝑡 𝑚𝑎𝑠𝑠 𝑜𝑓 𝑎𝑑𝑠𝑜𝑟𝑏𝑒𝑛𝑡

𝐶𝑒 = concentration of adsorbate remaining in the syrup after adsorption

𝐾1, 𝑞𝑜 are Langmuir constant representing the maximum adsorption capacity and energy of

adsorption respectively.

0.76

𝑞𝑒=

1

344.83 × 26.93+

0.76

344.83

0.76

𝑞𝑒= 2.3 × 10−3

𝑞𝑒 = 328.8 𝑚g/g Carbon (C)

Therefore 328.8 mg of adsorbate is adsorbed per every gram of adsorbent

Impurity Load = (𝐼𝑛𝑖𝑡𝑖𝑎𝑙 𝑐𝑜𝑛𝑐𝑒𝑛𝑡𝑟𝑎𝑖𝑜𝑛 − 𝑓𝑖𝑛𝑎𝑙 𝑐𝑜𝑛𝑐𝑒𝑛𝑡𝑟𝑎𝑡𝑖𝑜𝑛) 𝑚𝑔/𝑙 × 3314 𝐿/𝑑𝑎𝑦

Impurity Load = (7.6 − 0.76) 𝑚𝑔/𝑙 × 3314 𝐿/𝑑𝑎𝑦

= 22667 mg/day

200

BED PARAMETERS

Diameter of adsorbent Bed= 𝐷𝐵𝑒𝑑

𝐴𝐵𝑒𝑑 =𝜋𝐷2

𝐵𝑒𝑑

4

=𝜋 ×1.52

4

= 1.8 𝑚2

Calculating the volume of the adsorbent Bed

𝑉𝐵𝑒𝑑 = 𝐴Bed∙ HBed

For a bed height 𝐻𝐵𝑒𝑑 = 2 𝑚

Volume of bed, VBed =𝐴Bed∙ HBed

VBed = 1.8 ∙ 2

VBed = 3.6 𝑚3

The volume of the adsorbent bed to be used is 3.6 𝑚3

201

COLUMN DESIGN SPECIFICATIONS AND PARAMETERS

The volume of the column is the sum of the volume of the adsorbent bed (𝑉𝐵𝑒𝑑) and the volume

of the liquid-filled void(𝑉𝐿).

𝑉𝐴𝑑𝑠 = 𝑉𝐵𝑒𝑑 + 𝑉𝐿

The adsorbent should cover at least 66% of the volume of the adsorbent column

𝑉𝐵𝑒𝑑 = 66% ∗ 𝑉𝐴𝑑𝑠

VBed = 3.6 𝑚3

𝑉𝐴𝑑𝑠 = 𝑉𝐵𝑒𝑑

0.66=

3.6

0.66

𝑉𝐴𝑑𝑠 = 5.5 𝑚3

𝑉𝐿 = 𝑉𝐴𝑑𝑠 − 𝑉𝐵𝑒𝑑

𝑉𝐿 = 5.5 − 3.6

𝑉𝐿 = 1.9 𝑚3

The volume of the liquid-filled void, 𝑉𝐿is 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

𝐷𝐵𝑒𝑑 = 𝐷𝐴𝑑𝑠

𝑉𝐴𝑑𝑠 = 𝐴𝐴𝑑𝑠 ∗ 𝐻𝐴𝑑𝑠

𝐻𝐴𝑑𝑠 =4∗ 𝑉𝐴𝑑𝑠

𝜋𝐷2𝐵𝑒𝑑

202

𝐻𝐴𝑑𝑠 =5.5∗4

𝜋∗ 1.52

= 3 𝑚

BED POROSITY, 𝜺

The bed porosity is the void fraction of the reactor volume. It is given by the equation below

𝜀 =𝑉𝑜𝑖𝑑 𝑣𝑜𝑙𝑢𝑚𝑒(𝑉𝐿)

𝐴𝑑𝑠𝑜𝑟𝑏𝑒𝑟 𝑣𝑜𝑙𝑢𝑚𝑒(𝑉𝐴𝑑𝑠)=

𝑉𝐴𝑑𝑠−𝑉𝐵𝑒𝑑

𝑉𝐴𝑑𝑠= 1 −

𝑉𝐵𝑒𝑑

𝑉𝐴𝑑𝑠

= 1 −3.6

5.5

= 0.3

𝑇ℎ𝑒 𝑝𝑜𝑟𝑜𝑠𝑖𝑡𝑦 𝑜𝑓 𝑡ℎ𝑒 𝑏𝑒𝑑 , 𝜀 𝑖𝑠 0.3

FILTRATION RATE

The above volumetric flow of the glucose 3.98 𝑚3

𝑑𝑎𝑦 syrup will be pumped for 30 minutes that is

1809 seconds

𝑄𝐴𝑑𝑠 =3.98 𝑚3

1809 𝑠

= 0.0022 𝑚3

Therefore the filtration rate (FR) = 𝑄𝐴𝑑𝑠

𝐴𝐴𝑑𝑠

= 0.0022

1.8

= 0.001𝑚3

𝑚2𝑠

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, 𝜏.

Empty Bed Contact Time (EBCT)

This is the residence time for the empty adsorption column.

𝐸𝐵𝐶𝑇 =𝑉𝐴𝑑𝑠

𝑄

=5.5

0.0022

= 2500 𝑠

= 42 𝑚𝑖𝑛𝑢𝑡𝑒𝑠

Effective Contact Time, 𝝉

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.

𝜏 =𝜀∙𝑉𝐴𝑑𝑠

𝑄 = 𝐸𝐵𝐶𝑇 × 𝜀

= 42 × 0.3

= 12.5 𝑚𝑖𝑛

Therefore, the effective contact time, 𝜏 for a flow rate of 0.0022 𝑚3

𝑠 𝑖𝑠 12.5 𝑚𝑖𝑛utes

204

COSTING OF ADSORPTION COLUMN

The capital cost of the adsorption column takes into consideration

I. The cost of the column, C𝐴𝑑𝑠 given in Euros

II. The cost of adsorbent bed (fixed), C𝐵𝑒𝑑

THE COST OF THE COLUMN (𝐂𝑨𝒅𝒔)

The cost of the column is calculated using the following empirical equation

C𝐴𝑑𝑠 = 583.6 ∙ 𝐷0.675 ∙ 𝐻 ∙ 𝐹𝑚𝑎𝑡 ∙ (𝑃∙145

50)0.44

𝐷 = 𝑡ℎ𝑒 𝑑𝑖𝑎𝑚𝑒𝑡𝑒𝑟 𝑜𝑓 𝑐𝑜𝑙𝑢𝑚𝑛

𝐻 = 𝑡ℎ𝑒 ℎ𝑒𝑖𝑔ℎ𝑡 𝑜𝑓 𝑐𝑜𝑙𝑢𝑚𝑛

𝑝 = 𝑡ℎ𝑒 𝑤𝑜𝑟𝑘𝑖𝑛𝑔 𝑝𝑟𝑒𝑠𝑠𝑢𝑟𝑒

𝐹𝑚𝑎𝑡 = 𝑟𝑒𝑝𝑟𝑒𝑠𝑒𝑛𝑡𝑠 𝑎 𝑐𝑜𝑟𝑟𝑒𝑐𝑡𝑖𝑜𝑛 𝑡𝑜 𝑡𝑎𝑘𝑒 𝑖𝑛𝑡𝑜 𝑎𝑐𝑐𝑜𝑢𝑛𝑡 𝑓𝑜𝑟 𝑡ℎ𝑒 𝑐𝑜𝑠𝑡 𝑜𝑓 𝑡ℎ𝑒 𝑚𝑎𝑡𝑒𝑟𝑖𝑎𝑙

The parameters of the proposed design are

𝐷𝐴𝑑𝑠 = 1.5 𝑚

𝐻𝐴𝑑𝑠 = 3 𝑚

𝑃 = 100 𝐾𝑃𝑎

𝐹𝑚𝑎𝑡 𝑓𝑜𝑟 𝑠𝑡𝑎𝑖𝑛𝑙𝑒𝑠𝑠 𝑠𝑡𝑒𝑒𝑙 304 𝑖𝑠 1.7

C𝐴𝑑𝑠 = 583.6 ∙ 1.50.675 ∙ 3 ∙ 1.7 ∙ (120∙145

50)0.44

= € 51,385.6

C𝐴𝑑𝑠 = $ 57,551

205

THE COST OF THE FIXED BED(𝐂𝑩𝒆𝒅)

The cost of the packed or fixed bed is evaluated from the following equation

C𝐵𝑒𝑑 =𝜋∙𝐷2

4∙ 𝐻 ∙ 𝐶

Where

𝐻 = 𝑡ℎ𝑒 ℎ𝑒𝑖𝑔ℎ𝑡 𝑜𝑓 𝐵𝑒𝑑

D = 𝑡ℎ𝑒 ℎ𝑒𝑖𝑔ℎ𝑡 𝑜𝑓 𝐵𝑒𝑑

C= the volume cost of the packing (Bed). The cost of 1m3 of activated carbon is

3500 EUR which is about 3908.45 USD

𝐷𝐵𝑒𝑑=

𝐻𝐵𝑒𝑑

C𝐵𝑒𝑑 =𝜋∙𝐷2

4∙ 𝐻 ∙ 𝐶

=𝜋∙1.52

4∙ 2 ∙ 3500

= €12,372

C𝐵𝑒𝑑 = $13856

The total cost of the Adsorption column unit, CT = C𝐴𝑑𝑠 + C𝐵𝑒𝑑

=(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.

𝑉 = 𝜋𝑅2𝐻

𝑊ℎ𝑒𝑟𝑒 𝑉 = 𝑉𝑜𝑙𝑢𝑚𝑒 𝑜𝑓 𝑡ℎ𝑒 𝑒𝑛𝑡𝑖𝑟𝑒 𝑟𝑒𝑎𝑐𝑡𝑜𝑟

𝑅 = 𝑅𝑎𝑑𝑖𝑢𝑠 𝑜𝑓 𝑡ℎ𝑒 𝑟𝑒𝑎𝑐𝑡𝑜𝑟 = 0.9𝑚

𝐻 = 𝐻𝑒𝑖𝑔ℎ𝑡 𝑜𝑓 𝑡ℎ𝑒 𝑟𝑒𝑎𝑐𝑡𝑜𝑟 = 4.0𝑚

𝐷𝑖𝑎𝑚𝑒𝑡𝑒𝑟 𝑜𝑓 𝑡ℎ𝑒 𝑟𝑒𝑎𝑐𝑡𝑜𝑟, 𝐷 = 1.8𝑚

𝑇ℎ𝑒𝑟𝑒𝑓𝑜𝑟𝑒 𝑅𝑎𝑑𝑖𝑢𝑠, 𝑅 = 𝐷

2

𝑅𝑎𝑑𝑖𝑢𝑠 = 1.8

2 = 0.9𝑚

𝐻𝑒𝑛𝑐𝑒 𝑣𝑜𝑙𝑢𝑚𝑒 𝑜𝑓 𝑡ℎ𝑒 𝑡𝑎𝑛𝑘, 𝑉 = 𝜋 × (0.9)2 × 4.0

𝑉 = 10.2𝑚3

207

VOLUME OF REACTING CHAMBER (𝑽𝑹)

Still using the equation for the volume of a cylinder,

𝑉𝑅 = 𝜋 𝑟2ℎ

𝑊ℎ𝑒𝑟𝑒 𝑟 = 𝑅𝑎𝑑𝑖𝑢𝑠 𝑜𝑓 𝑡ℎ𝑒 𝑟𝑒𝑎𝑐𝑡𝑖𝑜𝑛 𝑐ℎ𝑎𝑚𝑏𝑒𝑟 = 0.80𝑚

ℎ = ℎ𝑒𝑖𝑔ℎ𝑡 𝑜𝑓 𝑡ℎ𝑒 𝑟𝑒𝑎𝑐𝑡𝑖𝑜𝑛 𝑐𝑎ℎ𝑚𝑏𝑒𝑟 = 3.8𝑚

𝐷𝑖𝑎𝑚𝑒𝑡𝑒𝑟 𝑜𝑓 𝑡ℎ𝑒 𝑟𝑒𝑎𝑐𝑡𝑖𝑜𝑛 𝑐ℎ𝑎𝑚𝑏𝑒𝑟 (𝑑) = 1.6𝑚

𝑅𝑎𝑑𝑖𝑢𝑠 𝑟 = 𝑑

2

𝑅𝑎𝑑𝑖𝑢𝑠 𝑟 = 1.6

2 = 0.80𝑚

𝑉𝑜𝑙𝑢𝑚𝑒 𝑜𝑓 𝑡ℎ𝑒 𝑟𝑒𝑎𝑐𝑡𝑖𝑜𝑛 𝑐ℎ𝑎𝑚𝑏𝑒𝑟 , 𝑉𝑅 = 𝜋 × (0.802) × 1.8

𝑉𝑅 ≈ 7.64𝑚3

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 <𝐷𝐼

𝐷𝑇< 12

Using similar conditions for the calculation of the diameter of the impellers, Assuming

𝐷𝐼

𝐷𝑇 = 0.3

𝑊ℎ𝑒𝑟𝑒 𝐷𝐼 = 𝐷𝑖𝑎𝑚𝑒𝑡𝑒𝑟 𝑜𝑓 𝑡ℎ𝑒 𝑖𝑚𝑝𝑒𝑙𝑙𝑒𝑟

208

𝐷𝑇 = 𝐷𝑖𝑎𝑚𝑒𝑡𝑒𝑟 𝑜𝑓 𝑡ℎ𝑒 𝑡𝑎𝑛𝑘 = 1.8𝑚

𝐷𝐼 = 0.3𝐷𝑇 = 0.3 × 1.8

𝐷𝐼 = 0.54 𝑚.

NUMBER OF IMPELLERS (N)

According to Davis (2009), the number of impellers N is given by

𝐻 − 𝐷𝐼

𝐷𝐼 > 𝑁 >

𝐻 − 2𝐷𝐼

2𝐷𝐼

𝑊ℎ𝑒𝑟𝑒 𝐻 = 𝑡ℎ𝑒 ℎ𝑒𝑖𝑔ℎ𝑡 𝑜𝑓 𝑡ℎ𝑒 𝑡𝑎𝑛𝑘 = 3.0 𝑚

𝑏𝑢𝑡 𝐻 − 2𝐷𝐼

2𝐷𝐼 =

3 − 2(0.42)

2(0.42)

𝑁 ≅ 3

𝑇ℎ𝑒𝑟𝑒𝑓𝑜𝑟𝑒 𝑡ℎ𝑒 𝑛𝑢𝑚𝑏𝑒𝑟 𝑜𝑓 𝑖𝑚𝑝𝑒𝑙𝑙𝑒𝑟𝑠 𝑖𝑠 3

IMPELLER TIP SPEED (n)

𝑛 = 𝜋 × 𝑁 × 𝐷𝑇

𝑛 = 𝜋 × 3 × 1.4

𝑛 = 13.19𝑚𝑠−1

𝑇ℎ𝑖𝑠 𝑖𝑠 𝑒𝑞𝑢𝑖𝑣𝑎𝑙𝑒𝑛𝑡 𝑡𝑜 180 𝑟𝑝𝑚

209

FLOW OF FLUID IN THE REACTOR

To identify the flow of the fluid in the reactor, the Reynolds number needs to be calculated.

𝑁𝑅 = 𝑑2𝑛𝜌

𝜇

𝑤ℎ𝑒𝑟𝑒 𝑑 = 𝑡ℎ𝑒 𝑑𝑖𝑎𝑚𝑒𝑡𝑒𝑟 𝑜𝑓 𝑡ℎ𝑒 ℎ𝑦𝑑𝑟𝑜𝑔𝑒𝑛𝑎𝑡𝑖𝑜𝑛 𝑟𝑒𝑎𝑐𝑡𝑜𝑟 = 1.8𝑚

𝑛 = 𝑡ℎ𝑒 𝑖𝑚𝑝𝑒𝑙𝑙𝑒𝑟 𝑡𝑖𝑝 𝑠𝑝𝑒𝑒𝑑 = 13.19𝑚𝑠−1

𝜌 = 𝑑𝑒𝑛𝑠𝑖𝑡𝑦 𝑜𝑓 𝑠𝑜𝑟𝑏𝑖𝑡𝑜𝑙 = 1490 𝑘𝑔𝑚−3

𝜇 = 𝑡ℎ𝑒 𝑣𝑖𝑠𝑐𝑜𝑠𝑖𝑡𝑦 𝑜𝑓 𝑠𝑜𝑟𝑏𝑖𝑡𝑜𝑙 𝑝𝑟𝑜𝑑𝑢𝑐𝑒𝑑 = 0.11 𝑃𝑎𝑠

𝑁𝑅 = (1.8)2 × 13.19 × 1490

0.11

𝑁𝑅 = 578873.12

𝑇ℎ𝑖𝑠 𝑖𝑠 𝑔𝑟𝑒𝑎𝑡𝑒𝑟 𝑡ℎ𝑎𝑛 10000 (𝑁𝑅

> 10000)𝑎𝑛𝑑 𝑡ℎ𝑒𝑟𝑒𝑓𝑜𝑟𝑒 𝑡ℎ𝑒 𝑓𝑙𝑜𝑤 𝑖𝑛𝑠𝑖𝑑𝑒 𝑡ℎ𝑒 𝑡𝑎𝑛𝑘 𝑖𝑠 𝑡𝑢𝑟𝑏𝑢𝑙𝑒𝑛𝑡.

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.

𝑝 = 𝑛𝑢𝑚𝑏𝑒𝑟 𝑜𝑓 𝑝𝑜𝑙𝑒𝑠 = 2

𝑓 = 𝑚𝑜𝑡𝑜𝑟 𝑜𝑝𝑒𝑟𝑎𝑡𝑖𝑛𝑔 𝑓𝑟𝑒𝑞𝑢𝑒𝑛𝑐𝑦

𝑛 = 𝑠𝑝𝑒𝑒𝑑 𝑜𝑓 𝑟𝑜𝑡𝑜𝑟 = 200 𝑟𝑝𝑚

210

𝑀𝑃 = 𝑚𝑜𝑡𝑜𝑟 𝑝𝑜𝑤𝑒𝑟 = 3 ℎ𝑜𝑟𝑠𝑒 𝑝𝑜𝑤𝑒𝑟 (𝐻𝑃)

SYNCHRONOUS SPEED (NS)

This refers to the rotation speed of the stators magnetic field.

𝑁𝑆 = 120 × 𝑓

𝑝

𝑁𝑆 = 120 × 10

5

𝑁𝑆 = 240 𝑟𝑝𝑚

TORQUE (T)

This is the force required to achieve one complete revolution

𝑇 = 𝑀𝑃 × 5252

𝑁𝑆

𝑇 = 3 × 5252

240

𝑇 = 65.65 𝐼𝑏𝑓𝑡

VOLUME OF CATALYST BED (V)

For the efficient conversion of glucose to sorbitol, the volumetric flow rate 𝑉𝑓 must be between

0.5- 3.5 volume of feed/ volume of catalyst.

But volume flow rate of feed of feed = 𝑉𝑓 = 0.2247 𝑚3/𝑚𝑖𝑛

𝑡ℎ𝑒 𝑣𝑜𝑢𝑚𝑒 𝑜𝑓 𝑓𝑒𝑒𝑑 𝑝𝑒𝑟 ℎ𝑜𝑢𝑟 = 13.482𝑚3/ℎ

𝑉𝑓

𝑉𝐶= 0.5

𝑉𝑜𝑢𝑚𝑒 𝑜𝑓 𝑐𝑎𝑡𝑎𝑙𝑦𝑠𝑡 = 0.5𝑉𝑓

211

𝑉𝑜𝑢𝑚𝑒 𝑜𝑓 𝑐𝑎𝑡𝑎𝑙𝑦𝑠𝑡 = 0.5 × 13.482 = 6.74 𝑚3

AREA OF CATALYST ON BED (A)

The height of the catalyst bed = 0.2m

𝐴𝑟𝑒𝑎 𝑜𝑓 𝑐𝑎𝑡𝑎𝑙𝑦𝑠𝑡 = 6.74

0.2= 33.7 𝑚2