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Transcript of saadurrehmansp11bec114d15ttt
PRODUCTION OF 2000 m3/hr OF SYN GAS FROM CO-GASIFICATION
USING CO2 SORBENT
SAAD UR REHMAN CIIT/SP11-BEC-114/LHR
M. RAHIL IFTIKHAR CIIT/SP11-BEC-098/LHR
SAJJAD AHMED CIIT/SP11-BEC-074/LHR
M. FAIZAN RAZZAQ CIIT/SP11-BEC-048/LHR
This report is submitted in partial fulfillment of
the requirements for the award of the degree of
Bachelor of Science in Chemical Engineering
Department of Chemical Engineering
COMSATS Institute of Information Technology
JANUARY 2
iii
To our Parents and Supervisor, Dr. Zakir Khan
xv
ACKNOWLEDGEMENTS
Firstly, we would like to thank ALLAH ALMIGHTY for providing us the strength and
the tolerance to successfully complete this project.
Next, we would like to thank our parents for providing us their moral support and
guiding us towards the right direction in approaching our task.
We would like to thank our head of department Dr. Asad U. Khan for providing us the
opportunity to complete this project.
We would like to express our great appreciation to Dr . Zaki r Khan for his valuable
and constructive suggestions during the planning and development of this project. His
willingness to give his time so generously has been very much appreciated.
We would also like to thank the supportive faculty of the department of chemical
engineering COMSATS Lahore. Finally, we would like to thank our class fellows for
their helpful suggestions during the course of the entire project.
xv
ABSTRACT
Gasification of carbonaceous solid fuels to convert into syngas for application of
power, liquid fuels and chemicals is practiced worldwide. Coal and biomass are major
feedstock for gasification. The concept of blending of coal and biomass will make the
gasification efficient and profitable. In this project mixture of coal and Sugarcane
Bagasse is used. Coal has high fixed carbon and ash content as compared to Sugarcane
Bagasse whereas Sugarcane Bagasse has high volatile matter as compared to coal.
Gasification with in situ carbon dioxide capture has good prospects for the production
of hydrogen enriched Synthesis gas. In this project, the mixture of coal and biomass is
dried and grinded before being fed to the gasifier. After gasification, the resultant gas
is cooled and cleaned by the help of Cyclone Separator and Scrubber. The resultant
gas has a lower heating value of 39630 KJ/Kg of Synthesis gas and consists of 76% of
Hydrogen on volume basis. There exists high potential for gasification production in
Pakistan due to presence of abundant coal sources and agriculture waste. Due to energy
and environmental issues, Synthesis gas has become a more attractive clean fuel.
Furthermore, gasification has become a sustainable technology for the production of
Synthesis gas.
xv
TABLE OF CONTENTS
CHAPTER TITLE PAGE
DECLARATION ii
DEDICATION iii
ACKNOWLEDGEMENT iv
ABSTRACT v
TABLE OF CONTENTS vi
LIST OF TABLES x
LIST OF FIGURES xi
LIST OF SYMBOLS xii
LIST OF APPENDICES xv
1 LITERATURE REVIEW 1
1.0 Coal 1
1.1. Types of coal 1
1.2. Coal Properties 4
1.2.1. Heating Value 4
1.2.2. Caking and Swelling Properties 4
1.2.3. Hardness 5
1.2.4. Density 5
1.3. Coal Reserves in Pakistan 6
1.4. Biomass 8
1.4.1. Types of biomass in Pakistan 8
1.4.2. Biomass Analysis 9
1.4.3. Production of Crops and their Residue availability in Pakistan 9
1.4.4. Properties of Feed Stock Suitable for Gasification 10
1.4.5. Selection of Feed Stock 13
1.5. Application of Syn Gas 14
1.5.1. Heat 14
1.5.2. Electricity 14
1.5.3. Combined heat and power 14
1.5.4. Transport fuel 15
xv
2 PROCESS SELECTION 16
2.1. Methods of Syn Gas Production 16
2.1.1. Landfill 16
2.1.2. Gasification 17
2.1.3. Pyrolysis 18
2.2. Gasifier Types 18
2.2.1. Counter-current fixed bed ("Up draft") gasifier 19
2.2.2. Co-current fixed bed ("down draft") gasifier 20
2.2.3. Fluidized bed reactor 21
2.2.4. Entrained Flow Gasifier 22
2.3. Gasifier Selection 23
3 PROCESS DESCRIPTION 24
3.1. Description of the Process 24
3.1.1. Feed Treatment 25
3.1.2. Gasification 25
3.1.3. Gas Cleaning 26
3.1.4. Utility Section 27
3.2. Factors Affecting Gasification 27
3.2.1. Temperature 27
3.2.2. Pressure 27
3.2.3. Fuel Bed Height 27
3.2.4. Fluidization Velocity 27
3.2.5. Feed Properties 28
3.3. Block Diagram 29
3.4. Process Flow Diagram 30
4 MATERIAL BALANCE 31
4.1. Material Balance on Dryer 31
4.2. Material Balance on Mill 32
4.3. Material Balance on Gasifier 32
4.3.1. Gasification Reactions 33
4.4. Material Balance on Cyclone Separator 36
4.5. Material Balance on Scrubber 37
4.6. Overall Material Balance 39
5 ENERGY BALANCE 40
5.1. Energy Balance on Dryer 40
5.1.1. Heat Inlet 40
5.1.2. Heat Outlet 41
5.2. Energy Balance on Ball Mill 42
5.3. Energy Balance on Gasifier 42
5.3.1. Heat Inlet 42
5.3.2. Heat involved in Chemical Reactions 43
5.3.3. Heat Outlet 45
xv
5.3.4. Overall Energy Balance 45
5.4. Energy Balance on Cyclone Separator 45
5.5. Energy Balance on Scrubber 46
5.6. Overall Energy Balance 48
6 EQUIPMENT DESIGN 49
6.1. Dryer Design 49
6.1.1. Brief Introduction 49
6.1.2. Classification of Dryers 50
6.1.3. Selection of dryer 50
6.1.4. Direct Heat Rotary Dryer 51
6.1.5. Dryer Design 52
6.2. Gasifier Design 62
6.2.1. Fluidized Bed Reactor 62
6.2.2. Hydro Dynamics 62
6.2.3. Minimum Fluidization Velocity 62
6.2.4. Diameter of Reactor 65
6.2.5. Height of Reactor 67
6.2.6. Volume of Reactor 68
6.2.7. Distributor Plate Design 68
6.2.8. Number of Orifices in the Distributor Plate 69
6.2.9. Gasifier Specification Sheet 70
6.2.10. Gasifier Design Data 70
6.3. Cyclone Separator Design 71
6.3.1. General design procedure 71
6.3.2. Number of Cyclones 71
6.3.3. Inlet Duct area 72
6.3.4. Dimensions of Cyclone 72
6.3.5. Calculation of Scaling Factor 73
6.3.6. Pressure Drop Calculation 74
6.3.7. Cyclone Separator Specification Sheet 75
6.3.8. Cyclone Separator Design Data 75
6.4. Design of Scrubber 76
6.4.1. Designing Steps 76
6.4.2. Calculation of Height of Transfer Units (Onda’s Method) 81
6.4.3. Calculation of Liquid Film Mass Transfer Coefficient 81
6.4.4. Calculation of Gas Film Mass Transfer Coefficient 82
6.4.5. Calculation of Gas-film Transfer Unit Height 82
6.4.6. Calculation of Liquid-film Transfer Unit Height 83
6.4.7. Calculation of Height of Transfer Units 83
6.4.8. Calculation of Height of Tower 83
6.4.9. Pressure Drop Calculations 84
6.4.10. Scrubber Specification Sheet 85
6.4.11. Scrubber Design Data 85
xv
7 INSTRUMENTATION AND PROCESS CONTROL 86
7.1. Instrumentation 86
7.2. The Concept of Measurement in Automation Application 86
7.3. Measurement 87
7.4. Process 87
7.4.1. Process Control 87
7.4.2. Objectives of Process Control 88
7.5. Basic Elements of Process Control 89
7.6. Basics of Process Control 90
7.7. Selection of Controller 90
7.8. Control Loops 92
7.9. Control Schemes of Gasifier 94
7.10. Control Schemes of Scrubber 94
8 COST ESTIMATION 96
8.1. Introduction 96
8.2. Fixed Capital Investment 97
8.3. Working Capital 97
8.4. Cost Index 97
8.5. Purchased Equipment Cost 98
8.5.1. Estimate of Dryer Cost 98
8.5.2. Estimate of Ball Mill Cost 98
8.5.3. Estimate of Gasifier Cost 99
8.5.4. Estimate of Cyclone Separator Cost 100
8.5.5. Estimate of Scrubber Cost 101
8.6. Estimation of Working Capital 104
8.7. Total Investments 104
8.8. Production Costs 104
9 HAZOP STUDY 107
9.1. Introduction 107
9.2. Objectives of Hazop Study 107
9.3. Keywords used in HAZOP STUDY 108
9.4. Primary Keywords 108
9.5. Secondary Keywords 109
9.6. How to Conduct a Hazop Study 109
9.7. HAZOP Method Flow Diagram 111
9.8. Hazop Analysis on Fluidized Bed Reactor 112
9.9. Hazop Analysis on Dryer: 113
REFERENCES 114
APPENDICES 117
xv
LIST OF TABLES
TABLE NO. TITLE PAGE
1.1 Particle and Bulk Density 5
1.2 Ultimate and Proximate Analysis of Biomass 9
1.3 Production of Crops and their Residue availability in Pakistan 10
3.1 Operating Conditions and Process Parameters in the Gasifier 26
4.1 Proximate analysis of Coal and Biomass (Dry Basis) 33
4.2 Gasification Reaction 34
4.3 Gasifier Material 35
4.4 Product Gas Composition of Gasifier 36
4.5 Composition of Gas at Inlet of Scrubber 37
4.6 Outlet composition of Gases of Scrubber 38
4.7 Material Balance on Gasifier 44
5.1 Composition of Inlet Gas of Scrubber 46
5.2 Composition of Outlet Gases of Scrubber 47
6.1 Temperature of Inlet and Outlet Streams 54
6.1 Bed Material Properties 63
6.2 Fluidizing Gas Properties (at 950K and 1 atm) 63
6.3 Fluidizing Gas Properties (at 523K and 2 atm) 63
6.4 Proximate analysis of Coal and Biomass (Dry Basis) 65
6.5 Scrubber Material Balance 76
6.6 Scrubber Top Composition 77
8.1 Total Purchased Cost of Major Equipments 102
8.2 Typical factors for the estimation of project fixed capital cost 103
8.3 Fixed Capital Cost 103
9.1 Primary Keywords 108
9.2 Secondary Keywords 109
xv
LIST OF FIGURES
FIGURE NO. TITLE PAGE
2.1 Counter-current Fixed bed (Up draft) Gasifier 19
2.2 Co-current Fixed bed (Down draft) Gasifier 20
2.3 Fluidized Bed Reactor 21
2.4 Entrained Flow Gasifier 22
3.1 Block Diagram 29
3.2 Process Flow Diagram 30
4.1 Overall Material Balance 39
5.1 Overall Energy Balance 48
6.1 Direct Heat Rotary Dryer 51
6.2 Hydrodynamics Calculations 62
6.3 High through-put Cyclone 71
7.1 Control Schemes of Gasifier 94
7.2 Control Schemes of Scrubber 95
9.1 HAZOP Method Flow Diagram 111
xv
LIST OF SYMBOLS
F Feed
LHV Lower Heating Value
HHV Higher Heating Value
CGR Char Gasification Reaction
SMR Steam Methane Reaction
WGSR Water Gas Shift Reaction
H Humidity
Xa Inlet moisture content
Xb Final moisture content
Ar Archimedes Number
Ƿf Fluid Density
Ƿp Density of the Particle
Ƿb Density of the Bed
Emf Voidage at minimum Fluidization
Øs Sphericity
Dp Particle Diameter
Umf Minimum Fluidization Velocity
Remf Renold Number at minimum fluidization
S/C Steam to Carbon Ratio
Qo Initial Volumatric Flow Rate
Qmf Volumatric Flow Rate at minimum Fluidization
Dbm Bubble Diameter
R Distributor to bed pressure ratio
Δpb Bed pressure drop
Δpd Distributor pressure drop
Cd Drag Coefficient
UOR Velocity at Orifice
NOR Number of Holes at Orifice
xv
Dc Diameter of Cyclone
TH Total Height
AO Outlet Area
DO Outlet Diameter
UT Terminal Velocity
D2/D1 Scaling Factor
Dc2 Diameter of Cyclone
Dc1 Diameter of Standard Cyclone=0.203 m
Q1 Standard Flow Rate=0.0619 m3/sec
Q2 Volumetric Flow Rate
ρ1 Standard solid-fluid density =2000 kg/m3
ρ2 Particle density
μ1 Standard viscosity =0.018 Ns/m2
μ2 Gas viscosity
FC Friction factor, taken as 0.005 for gases
AS Surface area of cyclone exposed to the spinning
AI Area of inlet Duct
RT Radius of circle to which the centre line of the inlet
RE Radius of exit pipe
Δp Pressure Drop
C Porosity of packing factor
a Surface area of packing (m2/m3)
FLV Flooding Ratio of Liquid to Vapor
A Column Area
VW Vapor Mass Flow Rate
HOG Height of Overall Gas-phase transfer units
HG Individual Film Transfer Units for Gas
NOG Number of Overall Gas-phase transfer units
HL Individual Film Transfer Units for Liquid
HOL Height of Overall Liquid-phase transfer units
xv
NOL Number of Overall Liquid-phase transfer units
Gm Molar Gas Flow-Rate per unit cross-sectional area
Lm Molar Liquid Flow-Rate per unit cross-sectional area
αW Effective interfacial area of packing per unit volume
α Actual area of packing per unit volume
δL Surface tension of liquid
DL Diffusivity of liquid
Ф Association factor
xv
LIST OF APPENDICES
APPENDIX TITLE PAGE
A1 Char Gasification Reaction Equilibrium 117
A2 Equilibrium Methane Conversion at S/C =2 117
A3 CO Conversion in Water Gas Shift Reaction 118
A4 Temperature Pattern in Dryer 118
A5 Cyclone Pressure Drop Factor 119
A6 Standard Cyclone Dimension, High Gas Rate Cyclone 119
A7 Performance Curves, High Gas Rate Cyclone 120
A8 Scaled Performance Curve Cyclone 120
A9 Generalized Pressure Drop Correlation 121
A10 Number of Transfer units NOG as a function of y1/y2 121
A11 Installed Cost of Dryer 122
A12 Columns Plates. Time Base mid-2004 122
CHAPTER 1
LITERATURE REVIEW
1.0. Coal
Coal is a fossil fuel formed from plants that were buried millions of years
ago. The high-temperature, high-pressure conditions underground transformed the
plants physically and chemically, forming coal. Coal contains energy that the plants
absorbed from the sun millions of years ago. Burning coal releases this energy [1].
1.1. Types of coal
As geological processes apply pressure to dead biotic material over time,
under suitable conditions it is transformed successively into
Peat
Lignite
Sub-bituminous Coal
Bituminous Coal
Anthracite
2
All coal has been formed from biomass. Over time this biomass has been
turned into peat. When covered under a layer of overburden, the influence of time,
pressure, and temperature converts this material into brown coal or lignite.
Subsequently, the later material will turn into sub-bituminous coal, then into
bituminous coal, and finally into anthracite. Coal is often classified in terms of its
rank, which increases from brown coal to anthracite. Brown coal, lignite, and sub-
bituminous coals are called low-rank coals, whereas higher-rank coals are often
called hard coals. The terms brown coal and lignite are essentially synonymous,
lignite being used more often in the United States and brown coal in Europe and
Australia.
1.1.1. Peat
Peat, considered to be a precursor of coal, has industrial importance as a
fuel in some regions, for example, Ireland and Finland. In its dehydrated form, peat
is a highly effective absorbent for fuel and oil spills on land and water. It is also
used as a conditioner for soil to make it more able to retain and slowly release
water.
1.1.2. Lignite
Lignite is a brownish-black fossil fuel that is used primarily for electric
power generation. Considered to be a low-ranking type of coal, the fuel is usually
categorized by geologists as a recent fuel. Typically, it fits between peat and sub-
bituminous coal on geological solid fuel ranking scales. Lignite coal is burnable
and may also be referred to as brown coal. Geologically, lignite is believed to be
relatively young in age. Some geologists estimate that it formed roughly 251
million years ago.
3
1.1.3. Sub-bituminous coal
Sub-bituminous coal, whose properties range from those of lignite to those
of bituminous coal, is used primarily as fuel for steam-electric power generation
and is an important source of light aromatic hydrocarbons for the chemical
synthesis industry.
1.1.4. Bituminous coal
Bituminous coal or black coal is a relatively soft coal containing a tarlike
substance called bitumen. It is of higher quality than lignite coal but of poorer
quality than anthracite coal. Bituminous coal is an organic sedimentary rock
formed by diagenetic and submetamorphic compression of peat bog material. The
carbon content of bituminous coal is around 60-80%; the rest is composed of water,
air, hydrogen, and sulfur. [2].
1.1.5. Anthracite
Anthracite coal is a form of coal that is almost made entirely of carbon.
Anthracite coal is much harder than other forms of coal such as bituminous, and is
usually found in areas surrounding mountains or deep valleys. Anthracite burns
much cleaner than other forms of coal due to its low pollutant content. In fact,
anthracite may contain 91% to 98% pure carbon, leaving only 2% to 9% of other
elements. Anthracite coal is difficult to ignite and burns with a blue, smokeless
flame [3].
4
1.2. Coal Properties
1.2.1. Heating Value
The heating value is obtained by combustion of the sample in a calorimeter.
If not available, the heating value can be calculated with, for example, the Dulong
formula [4] from the ultimate analysis:
HHV (Btu/lb) = 14,600 C + 62,000 (H –O/8 ) + 4000 S
Where C, H, O, and S are the mass fractions of the elements obtained from
the ultimate analysis. There are other formulae for calculating the heating value
from the ultimate and /or proximate analysis. It is always useful to calculate the
heating value from these analyses, as it is a good cross check on measured values.
If the deviation is more than a few percent, all analyses must be checked.
1.2.2. Caking and Swelling Properties
Another important property of a coal is the swelling index. The swelling
index is determined by heating a defined sample of coal for a specified time and
temperature, and comparing the size and shape taken by the sample with a defined
scale. There are a number of different scales defined in, for example, ASTM D
720-91, BS 1016, or ISO 335. The swelling index is an indicator for the caking
properties of a coal and its expansion on heating. Softening/caking does not occur
at a certain temperature but over a temperature range. It is an important variable
for moving-bed and fluid-bed gasifier. For the gasifier of entrained-flow systems,
the coal softening point has no relevance. However, the softening point may limit
the amount of preheating of the pulverized coal feedstock used in dry coal feed
gasifier.
5
1.2.3. Hardness
Physical properties are not very relevant for the operation of a gasifier as
such. The hardness of the coal is, for example, mainly important for the milling and
grinding up stream of the gasifier. The hardness of a coal is usually dependant on
the nature and quantity of its ash content, although some coals, such anthracites,
are also hard. High ash content or a very high hardness of the ash in the coal can
make a feedstock unattractive for gasification because of the high cost of milling
and grinding. Ashes with high silica and/or alumina contents have a high hardness.
The hardness is generally characterized by the hard grove grind ability index.
1.2.4. Density
The density is primarily of importance for the transport of the coal. In this
connection, it is important to discriminate between the particle density and the bulk
density of the coal. The bulk density is always lower, as is shown in table 1.1:
Table 1.1: Particle and Bulk Density [5]
Fuel Particle (True) Density
(kg/m3)
Bulk (Apparent)
Density (kg/m3)
Anthracite 1450-1700 800-930
Bituminous coal 1250-1450 670-910
Lignite 1100-1250 550-630
6
1.3. Coal Reserves in Pakistan [6]
Coal Reserves of Pakistan and Azad Kashmir (million tons).
Coal Reserves of Balochistan Province (million tons). Only Chamalang, Khost-Shahrig-Harnai, Duki-
anambar, Mach-AbeGum, and Pir Ismail Ziarat-Marwar coalfields are developed so far.
Coalfield Coal Th. M. Moist. V.M. FixCarbon Ash T.
Sulphur
H.V.; BTU/lb Rank
Chamala
ng
0.2-2.5m 6 1.14-4.58 7.68-48.17 0.65-50.05 5.35-84.96 3.44-6.93 1818-13569 LigC to hvBb
Kingri
(K)
0.5-2.5m 3.9 1.64 18.4 25.1 55.2 5.58 2000-10,000 LigC to
SubC
Kingri-hikar
0.2-2.0m 1 Same as Chamalang coalfields
Narwal-
Dab
0.2-2.0m 1 Same as Chamalang coalfields
Toi
Nala/
Ghoze
Ghar
0.3-2.0m 1.2 1.8-1.9 42.3-42.9 32.1-32.9 22.8-23.1 5.8-6.1 9,790-13,000 SubC to
hvBb Khost-
Sharig-
Harnai
0.3-2.3m 20.9 1.7-11.2 9.3-45.3 25.5-43.8 9.3-34.0 3.5-9.5 9,637-15,499 SubC to
hvCb Sor Range- Deghari
0.3-1.3m 9.8 3.9-18.9 20.7-37.5 41.0-50.8 4.9-17.2 0.6-5.5 11,245-13,900 SubA to hvBb
Duki-
Anam
bar
0.2-2.3m 22.8 3.5.11.5 32.0-50.0 28.0-42.0 5.0-38.0 4.0-6.0 10,131-14,164 SubB to
hvAb Mach-
Abegum
0.6-1.3m 9 7.1-12.0 34.2-43.0 32.4-41.5 9.6-20.3 3.2-7.4 11,110-12,937 SubA to
hvCb Pir Ismail
Ziarat
0.4-0.7m 3.6 6.3-13.2 34.6-41.0 19.3-42.5 10.3-37.5 3.2-7.4 10,131-14,164 SubB to
hvAb Johan 0.1-0.3m 0.25 Same as Mach coalfield
Total 79.45
Abbreviations; Coal Th - Cumulative coal thickness, M-measured, Ind-indicated, Hyp-hypothetical, Moist- moisture, T.
sulphur-total sulphur, H.V.-heating value, BTU/lb - British thermal unit/pound, m- metre, lig- lignite, Sub-sub
bituminous, b-bituminous, hv-high volatile.
Coal Reserves of Punjab (million tons). Both coalfields are developed.
Province Measured Indicated Inferred Hypothetical Total Reserves
Sindh 3339 11835 56646 113637 185457
Balochistan 79.45 150.45 183.5 45.3 458.7
Punjab 57 31 2 145 235
Khyber Pakhtunkhwa 3 5.75 109.24 5 122.99
Azad Kashmir 1 1 6.72 - 8.72
Grand Total 3479.45 12023.20 56947.26 113832.30 186282.41
Coalfield Coal Th. M. Moist V.M. FixCarbon Ash T. Sulphur H.V.; BTU/lb Rank
Makerwal
0.3-2.0m
7
2.8-6.0
31.5-48.1
34.9-44.9
6.4-30.8
2.8-6.3
10,688-14,029
SubA to hvAb
Salt Range
0.15-1.2m
50
3.2-10.8
21.5-38.8
25.7-44.8
12.3-44.2
2.6-10.7
10,131-14,164
SubC to hvAb
Total 57
7
Coal Reserves of Sindh Province (million tons). Only Lakhra and Meting-Jhimpir coalfields are
developed.
Coal Reserves of Khyber Pakhtunkhwa Province (million tons). Hangu/Orakzai, Cherat, Gulakhel
and Dara Adamkhel coalfields are developed so far.
Coalfield Coal Th. M. Moist. V.M. Fix Carbon Ash T.Sulphur H.V.; BTU/lb Rank
Hangu 0.4-0.6m 2 0.2-2.5 16.2-33.4 21.8-49.8 5.3-43.3 1.5-9.5 10,500-14,149 SubA- hvAb
Cherat 0.8-1.2m 0.5 0.1-7.1 14.0-31.2 37.0-76.9 6.1-39.0 1.1-3.5 9,386-14,171 SubC- hvAb
Gulakhel 0.3-2.0m - 2.8-6.0 31.5-48.1 34.9-44.9 6.4-30.8 2.8-6.3 10,688-14,029 SubA to hvAb
Shirani 0.1-0.3m 0.25 Same as Toi Nala coalfield
Dara
Adamkhe
l
0.4-0.6 0.25 Same as Hangu / Cherat coalfields
Total 3.00
Coal Reserves of Azad Kashmir (million tonnes). Kotli coalfield is developed so far.
Coalfield Coal
Th.
M. Moist. V.M. Fix Carbon Ash T. Sulphur H.V.; BTU/lb Rank
Kotli 0.2-
1.0m
1 0.2-6.0 5.1-32.0 26.3-69.5 3.3-50.0 0.3-4.8 7,336-12,338 LigA-hvCb
Total 0.2-
1.0m
1
Coalfield Coal Th. M. Moist. V.M. Fix
Carbon
Ash T.Sulphur H.V.; BTU/lb Rank
Lakhra 0.3-3.3m 244 9.7-38.1 18.3-38.6 9.8-38.2 4.3-49 1.2-14.8 5,503-9,158 LigB-SubC
Meting
-
Jhimpi
r
0.3-1.0m 10 26.6-36.6 25.2-34.0 24.1-32.2 8.2-16.8 2.9-5.1 7,734-8,612 LigA-SubC
Sonda- Thatta 0.3-1.5m 60 22.6-48.0 16.1-36.9 8.9-31.6 2.7-52.0 0.2-15.0 8,878-13,555 SubC-hvBb
Jherruck 0.3-6.2m 106 9.0-39.5 20.0-44.2 15.0-58.8 5.0-39.0 0.4-7.7 8,800-12,846 SubC-hvCb
Ongar 0.3-1.5m 18 9.0-39.5 20.0-44.2 15.0-58.8 5.0-39.0 0.4-7.7 5,219-11.172 LigB-SubA
Indus
zast
0.3-2.5m 51 9.0-39.5 20.0-44.2 15.0-58.8 5.0-39.0 0.4-7.7 7,782-8,660 LigA-SubC
Badin 0.5-3.1m 150 9.0-39.5 20.0-44.2 15.0-58.8 5.0-39.0 0.4-7.7 11,415-11,521 LigB-SubA
Thar 0.2-22.8m 2700 29.6-55.5 23.1-36.6 14.2-34.0 2.9-11.5 0.4-2.9 6,244-11,045 LigB-SubA
Total 3339
8
1.4. Biomass
Biomass refers to any organic material derived from plants that use sunlight
to grow. When burned, the energy stored in biomass is released to produce heat or
electricity. Common forms of solid biomass include agricultural crops, crop
residues and forestry products. Using biomass for energy offers potential
advantages:
Biomass is an abundant and renewable source of energy.
Using biomass for energy would diversify the energy supply
and reduce dependency on fossil fuels.
Biomass production may create new jobs for the local
economy.
1.4.1. Types of biomass in Pakistan
Following types of biomass is available in Pakistan.
1. Woody Biomass
2. Crop Residues
I. Rice Straw
II. Rice Husk
III. Rice Stalk
IV. Cotton Stalk
V. Cotton Husk
VI. Wheat Stalk
VII. Sugarcane Bagasse
VIII. Maize Cobs
IX. Maize Stalk
X. Barley Stalk
3. Animal Dung
I. Cattle
9
1.4.2. Biomass Analysis [7]
The Ultimate and proximate analysis of different type of biomass present in
Pakistan is shown in the following table.
Table 1.2: Ultimate and Proximate Analysis of Biomass
Type Of
Biomass
Proximate analysis Ultimate analysis CV
MJ/Kg Volatile
matter
%
FC
%
Ash
%
C
%
H
%
O
%
N
%
S
%
Wood [13] 83.0 7.2 9.8 50.5 6.1 43 0.3 0.1 18.7
Rice Straw [8] 65.5 15.8 18.7 38.2 5.20 36.26 0.87 0.18 15.09
Rice Husk [12] 61.81 16.95 21.24 38.50 5.20 34.61 0.45 14.7
Rice Stalk [13] 63.52 16.22 20.26 39.77 5.53 53.64 0.82 0.24 17.62
Cotton Stalk [10] 73.10 22 4.9 45.2 4.4 14.5 1.0 0.0 17.7
Cotton Husk [15] 72.8 20.59 6.61 47.03 5.96 38.42 1.8 16.9
Wheat Stalk [9] 75.3 17.7 7.0 43.4 6.0 44.5 0.8 0.1 18
Sugarcane
Bagasse [10]
71.9 18.03 10.07 45.58 5.96 45.21 0.15 0.001 18.7
Maize Cobs [12] 87.4 11.5 1.1 49 5.4 44.20 0.4 0.0 16.8
Maize Stalk [12] 73.15 19.2 7.65 44.73 5.87 40.44 0.60 0.07 17.7
Barley Stalk [11] 68.8 20.9 10.3 39.92 5.27 43.81 1.25 - 16.6
Cattle Dung [16] 71.24 15.4 13.38 42.07 5.60 50.0 1.75 17.0
1.4.3. Production of Different Crops and their respective Residue
availability in Pakistan [7]
The production of different crops in Pakistan with respect to their residue
availability is shown in the following table.
10
Table 1.3: Production of Different Crops and their respective Residue availability
in Pakistan
1.4.4. Properties of Feed Stock Suitable for Gasification
The composition of different materials varies significantly. Fuel
performance is related to the composition of the material. Important factors include
Moisture, ash, carbon, hydrogen, nitrogen, sulphur, oxygen and chloride content.
I. Moisture
Moisture content is the key factor determining the net energy content of
biomass material. Dry biomass has a greater heating value (or net energy potential),
as it uses little of its energy to evaporate any moisture. Increased moisture means
less energy available for the boiler.
Moisture content is usually reported on an “as-is” or wet basis in which the
water content is given as a fraction of the total weight. All biomass materials
contain some moisture, from as low as 8% for dried straw to over 50% for fresh-
Name of the crop Annual
production
(thousand MT)
Type of
residue
Crop to residue ratio
(residue/kg crop)
Total available
residue (thousand
MT)
Rice 6883 Husks 0.2 1376.6
Stalks 1.5 10,324.5
Straw 1.5 10,324.5
Cotton 3000 Husk 1.1 3300
Stalks 3.8a 11,802.8
Wheat 23,864 Stalks 1.5 35796
Sugarcane 49,373 Bagasse 0.33 16,293.09
Maize 296 Cobs 0.3 88.8
Stalks 2 592
Barley 82 Stalks 1.3 106.6
Dry chilly 187.7 Stalks 1.5 281.55
11
cut wood. A high moisture content adversely affects the collection, storage, pre-
processing, handling and transportation of biomass. In addition, transporting wet
material costs more.
Moisture content
Increases heat loss, due to evaporation and superheating of vapor
Helps, to a limit, in binding fines.
Aids radiation heat transfer.
The moisture content of raw biomass can be reduced by:
Leaving biomass in the field to dry for several weeks
Storing biomass, sheltered from precipitation
Commercial drying
II. Fixed carbon
Fixed carbon is the solid fuel left in the furnace after volatile matter is
distilled off. It consists mostly of carbon but also contains some hydrogen, oxygen,
sulphur and nitrogen not driven off with the gases. Fixed carbon gives a rough
estimate of heating value of coal.
III. Volatile Matter
Volatile matters are the methane, hydrocarbons, hydrogen and carbon
monoxide, and incombustible gases like carbon dioxide and nitrogen found in coal.
Thus the volatile matter is an index of the gaseous fuels present. Typical range of
volatile matter is 20 to 35%.
Volatile Matter content
Proportionately increases flame length, and helps in easier ignition of coal.
Influences secondary air requirement.
Influences secondary oil support.
12
IV. Ash
The non-combustible content of biomass is referred to as ash. High ash
content leads to fouling problems. Unfortunately biomass fuels, especially
agricultural crops/residues tend to have a high ash with high potassium content. As
a result, the ash melts at lower temperatures, resulting in “clinkers” that can jam
furnace elements. Alternately, slagging and fouling occur when ash is vaporized
and condensed in the boiler, resulting in the production of hard formations on the
heat transfer surfaces.
Wood has less than 3% ash. Agricultural crops have higher ash content,
from 3% and higher. Some boilers/stoves cannot handle fuels with high ash
content. More ash means more maintenance. [9]
Ash content:
Reduces handling and burning capacity.
Increases handling costs.
Affects combustion efficiency and boiler efficiency
Causes clinkering and slagging.
V. Carbon
The carbon content of biomass is around 40%, while coal contains 60% or
greater (Demirbas, 2007). A higher carbon content leads to a higher heating
value. [9]
VI. Hydrogen
The hydrogen content of biomass is around 6% (Jenkins, 1998). A higher
hydrogen content leads to a higher heating value. [9]
13
VII. Nitrogen
The nitrogen content of biomass varies from 0.2% to more than 1%
(Jenkins, 1998). Fuel bound nitrogen is responsible for most nitrogen oxide (NOx)
emissions produced from biomass combustion. Lower nitrogen content in the fuel
should lead to lower NOx emissions. [9]
VIII. Sulphur
Most biomass fuels have a sulphur content below 0.2%, with a few
exceptions as high as 0.5%–0.7%.Coals range from 0.5%–7.5% (Demirbas 2007).
Sulphur oxides (SOx) are formed during combustion and contribute significantly to
particulate matter (PM) pollution and acid rain. [9]
Since biomass has negligible sulphur content, its combustion does not
contribute significantly to sulphur emissions. It:
Affects clinkering and slagging tendencies
Corrodes chimney and other equipment such as air heaters and
economizers
Limits exit flue gas temperature.
1.4.5. Selection of Feed Stock
The selection of an appropriate feed stock is the important step at the initial
stage of gasification. The various samples are analyzed and compared in terms of
their costs and compositions. The comparison is generally made by the percentages
of sulfur content, fixed carbon, oxygen, ash and other volatile content. The Thar
coal and Sugarcane bagasse are selected as a feedstock for Gasification because of
i. High Calorific value
ii. Low Sulphur content
iii. High availability in Pakistan
14
1.5. Application of Syn Gas
Syngas can be used for heat production and for generation of mechanical
and electrical power. Like other gaseous fuels, producer gas gives greater control
over power levels when compared to solid fuels, leading to more efficient and
cleaner operation.
1.5.1. Heat
Gasifiers offer a flexible option for thermal applications, as they can be
retrofitted into existing gas fueled devices such as ovens, furnaces, boilers, etc.
where syngas may replace fossil fuels. Heating values of syngas are generally
around 4-10 MJ/m3.
1.5.2. Electricity
Currently Industrial-scale gasification is primarily used to produce
electricity from fossil fuels such as coal, where the syngas is burned in a gas
turbine. Gasification is also used industrially in the production of electricity,
ammonia and liquid fuels (oil) using Integrated Gasification Combined Cycles
(IGCC), with the possibility of producing methane and hydrogen for fuel cells.
IGCC is also a more efficient method of CO2 capture as compared to conventional
technologies. IGCC demonstration plants have been operating since the early
1970s and some of the plants constructed in the 1990s are now entering commercial
service.
1.5.3. Combined heat and power
In small business and building applications, where the wood source is
sustainable, 250-1000 kW and new zero carbon biomass gasification plants have
been installed in Europe that produce tar free syngas from wood and burn it in
reciprocating engines connected to a generator with heat recovery. This type of
plant is often referred to as a wood biomass CHP unit but is a plant with seven
15
different processes: biomass processing, fuel delivery, gasification, gas cleaning,
waste disposal, electricity generation and heat recovery.
1.5.4. Transport fuel
Diesel engines can be operated on dual fuel mode using producer gas. Diesel
substitution of over 80% at high loads and 70-80% under normal load variations
can easily be achieved. Spark ignition engines and fuel cells can operate on 100%
gasification gas. Mechanical energy from the engines may be used for e.g. driving
water pumps for irrigation or for coupling with an alternator for electrical power
generation.
CHAPTER 2
PROCESS SELECTION
2.1. Methods of Syn Gas Production
Different kinds of processes are involved for production of syngas. They
are
Landfill
Gasification
Pyrolysis
2.1.1. Landfill
The placement of solid waste in landfills is the probably the oldest and
definitely the most prevalent form of ultimate garbage disposal. From the outset, it
must be recognized that many landfills are nothing more than open, sometimes
controlled, dumps. The difference between landfills and dumps is the level of
engineering, planning, and administration involved.
17
Open dumps are characterized by the lack of engineering measures, no leach
ate management, no consideration of landfill gas management, and few, if any,
operational measures such as registration of users, control of the number of tipping
fronts or compaction of waste.
2.1.2. Gasification
Gasification is a process that converts organic or fossil based carbonaceous
materials into carbon monoxide, hydrogen and carbon dioxide. This is achieved by
reacting the material at high temperatures, without combustion, with a controlled
amount of oxygen and/or steam. The resulting gas mixture is called syngas (from
synthesis gas or synthetic gas) or producer gas and is itself a fuel.
The power derived from gasification of biomass and combustion of the
resultant gas is considered to be a source of renewable energy; the gasification of
fossil fuel derived materials such as plastic is not considered to be renewable
energy.
Gasification Reactions
Combustion gases can be produced by the reaction of the coal, char or
volatile matter with oxygen, carbon dioxide, hydrogen or stream. The main
reactions are listed below (for simplicity, only reactions with carbon are shown).
Partial combustion reaction: C + ½ O2 CO
Boudouard reaction: C+CO2 2CO
Hydro gasification reaction C + 2H2 CH4
Char Gasification reaction: C + H2O CO + H2
Combustion reaction C+O2 CO2
Water Gas Shift reaction CO + H2O CO2 + H2
18
Methane Reforming Reaction CH4 + H2O CO + 3H2
Carbonation Reaction CaO + CO2 CaCO3
2.1.3. Pyrolysis
Pyrolysis is a thermo chemical decomposition of organic material at
elevated temperatures without the participation of oxygen. It involves the
simultaneous change of chemical composition and physical phase, and is
irreversible. The word is coined from the Greek-derived elements pyr "fire" and
lysis "separating".
Pyrolysis is a case of thermolysis, and is most commonly used for organic
materials, being, therefore, one of the processes involved in charring. The pyrolysis
of wood, which starts at 200–300 °C (390–570 °F), occurs for example in fires
where solid fuels are burning or when vegetation comes into contact with lava in
volcanic eruptions. In general, pyrolysis of organic substances produces gas and
liquid products and leaves a solid residue richer in carbon content, char. Extreme
pyrolysis, which leaves mostly carbon as the residue, is called carbonization.
2.2. Gasifier Types
Depending upon the gasification medium, gasifiers can be classified into three
groups:
1. Air-blown, where air is the gasification medium
2. Oxygen-blown, where pure oxygen is the gasification medium
3. Steam Blown, where steam is the gasification medium
Air gasification produces a low heating value (5000 to 6000 kJ/kg or 3 to 6
MJ/m3, LHV) gas, which contains about 50% nitrogen and can fuel engines and
furnaces. Oxygen blowing is free from diluents like nitrogen. As a result it produces
higher (15,000 kJ/kg or 10 to 12 MJ/m3, LHV) heating value gas, Steam blown
gasifiers produce Syn gas with high hydrogen content and typically having a
heating value of 50,000 kJ/kg.
19
Depending upon how the gas and fuel contact each other, gasifiers can be
further divided into following types:
2.2.1. Counter-current Fixed bed (Up draft) Gasifier
A fixed bed of carbonaceous fuel (e.g. coal or biomass) through which the
"gasification agent" (steam, oxygen and/or air) flows in counter-current
configuration. The ash is either removed in the dry condition or as a slag. The
slagging gasifiers have a lower ratio of steam to carbon. Achieving temperatures
higher than the ash fusion temperature. The nature of the gasifier means that the
fuel must have high mechanical strength and must ideally be non-caking so that it
will form a permeable bed, although recent developments have reduced these
restrictions to some extent. The throughput for this type of gasifier is relatively low.
Thermal efficiency is high as the temperatures in the gas exit are relatively low.
However, this means that tar and methane production is significant at
typical operation temperatures, so product gas must be extensively cleaned before
use. The tar can be recycled to the reactor
Figure 2.1: Counter-current fixed bed (Up draft) gasifier [17]
20
2.2.2. Co-current Fixed bed (Down draft) Gasifier
Similar to the counter-current type, but the gasification agent gas flows in
co-current configuration with the fuel (downwards, hence the name "down draft
gasifier"). Heat needs to be added to the upper part of the bed, either by combusting
small amounts of the fuel or from external heat sources. The produced gas leaves
the gasifier at a high temperature, and most of this heat is often transferred to the
gasification agent added in the top of the bed, resulting in an energy efficiency on
level with the counter-current type. Since all tars must pass through a hot bed of
char in this configuration, tar levels are much lower than the counter-current type.
Figure 2.2: Co-current fixed bed ("down draft") gasifier [17]
21
2.2.3. Fluidized bed reactor
The fuel is fluidized by oxygen and steam or air. The temperatures are relatively
low in fluidized bed gasifiers, so the fuel must be highly reactive; low-grade coals
are particularly suitable. The agglomerating gasifiers have slightly higher
temperatures, and are suitable for higher rank coals. Fuel throughput is higher
than for the fixed bed, but not as high as for the entrained flow gasifier. The
conversion efficiency is higher than fixed bed. Recycle or subsequent combustion
of solids can be used to increase conversion. Fluidized bed gasifiers offers
improved mass and heat transfer from fuel, higher heating value and higher
efficiency.
The concern for climate change has increased the interest in biomass
gasification for which fluidized bed gasifiers are particularly popular, occupying
nearly 20% of their market. Fluidized bed gasifiers are divided into the following
two major types.
1. Bubbling fluidized bed gasifier
2. Circulating fluidized bed gasifiers
Figure 2.3: Fluidized Bed Reactor [17]
22
2.2.4. Entrained Flow Gasifier
Entrained-flow systems gasify pulverized fuel particles suspended in a
stream of oxygen (or air) and steam. Ash in the coal melts at the high operating
temperature of the gasifier and is removed as liquid slag. A number of
manufacturers offer commercial entrained-bed gasifiers for large-scale
applications, such as Texaco, Shell, and Kopper–Totzek. These gasifiers typically
operate at pressures up to 35 bar and use oxygen as the gasifier medium (William
et al., 2000). Many IGCC plants utilize entrained bed gasifiers.
Entrained bed gasifiers are available in much larger capacities (.100 MWe)
than other types, but these are more commonly used for fossil fuels like coal,
refinery wastes, etc. Their use for gasification is rather limited, as it requires the
fuel particles to be very fine (in the order of 80 to 100 mm).
In entrained flow gasifiers, Gasification occurs at 1600 ºC and the product
gas is taken out through collector pipe at 1000-1300 ºC. Ash is removed from the
bottom in the form of slag.
Figure 2.4: Entrained Flow Gasifier [19]
23
2.3. Gasifier Selection
Fluidized bed gasifier is selected due to its following advantages over other
type of gasifiers.
Short residence time
Uniform temperature distribution and Uniform Particle mixing.
Low char or/and tar contents
High cold gas energy efficiency
Reduced ash-related problems
Low air pollution.
No production of SOx because SO2, SO3 etc. are captured by lime stone.
Higher heating value of Synthesis gas
CHAPTER 3
PROCESS DESCRIPTION
3.1. Description of the Process
A simplified process has been developed for Synthesis gas production from
biomass and coal using pure steam as gasification agent and CaO as CO2 sorbent.
The block diagram of the process is shown. The whole process is consists of four
sections
1. Feed treatment
2. Gasification
3. Gas cleaning section
4. Utility section
The detail of each section is described in next headings. The process flow
diagram (PFD) is shown at the end of this chapter
25
3.1.1. Feed Treatment
Pretreatment of feedstock for gasifier is generally consisting of drying and
size reduction. Drying used to remove the moisture from the coal and biomass
either from flue gases or by air. Usually drying removes the moisture contents from
10-15 % in the biomass.
The best condition of feedstock for fluidized bed gasifier is that the
feedstock must well grind as well. So to achieve such best condition for coal and
biomass feed to gasifier a rotary dryer and ball mill used to remove moisture from
the feedstock and fine grinding respectively shown in PFD. [20]
3.1.2. Gasification
The conversion of feedstock to the Synthesis gas takes place in fluidized
bed gasifier through steam gasification process integrated with CO2 capture. There
are few assumptions were considered in flowsheet development modeling for
gasification process are as follows.
1. The gasifier operates under steady state conditions and atmospheric
pressure.
2. The reactions proceed adiabatically and at constant volume.
3. There is no tar formation in this process.
In the modeling framework, it is assumed that four reactions take place
simultaneously including Char gasification, Methane reforming, Water gas shift
and Carbonation. [22]
Char Gasification reaction
C(s) + H2O CO + H2
Steam Methane Reforming
CH4 + H2O CO + 3H2
Water Gas shift reaction
CO + H2O CO2 + H2
26
Carbonation reaction
CaO + CO2 CaCO3
Table 3.1: Operating Conditions and Process Parameters in the Gasifier
Parameter Value Reason
Temperature 950 K High H2/CO ratio [21]
Pressure 1 atm High H2/CO ratio [20]
Bed Material CaO As a CO2 sorbent [20] [21]
Steam/feed ratio 2 For hydrogen purity. [20]
Sorbent/feed ratio 1 For both hydrogen purity and
hydrogen yield profiles.[20]
Feed Size 2 mm For high fluidization [23]
3.1.3. Gas Cleaning
The product gas produced by the gasification process contained hydrogen,
carbon monoxide, carbon dioxide, methane, steam and fly ash. To get pure
hydrogen as end product, there are several steps involved in product gas cleaning
with different units like Cyclone Seprator and scrubber as shown in PFD. Fly ash
was removed from the system by Cyclone and the steam was removed by passing
through scrubber with fresh water. Along with the steam there are also some others
product gases will be also absorbs in water which was calculated by chart of
solubility of gases in water at atmospheric pressure and different temperature. The
scrubber is also used to cool down the product gas and to remove the remaining
amount of the solid particles. [20]
27
3.1.4. Utility Section
I. Steam Generation
The process design includes a steam generation system that produced steam
by a steam generator/Boiler. Furthermore steam is superheated to 523 K. The steam
is supplied to the gasifier at 2 atm pressure which reduces to 1 atm in the gasifier
due to energy losses. The steam generation system is also shown in PFD.
3.2. Factors Affecting Gasification
The main factors affecting the Gasification are as follows.
3.2.1. Temperature
Increasing the temperature increases the formation of combustible gases, decreases
the yield of char and liquids and leads to more complete conversion of the fuel. The
energy content of the syngas increases steadily up to 700 ºC then decreases at
higher temperatures.
3.2.2. Pressure
The rate of char gasification increase with increasing pressure, and the impacts are
most significant at high temperatures.
3.2.3. Fuel Bed Height
For a given reactor temperature, higher fuel bed heights increase the time fuels are
available for reactions to occur (residence time), which increases total syngas
yields and increases the concentrations of hydrogen, and carbon monoxide in the
syngas.
3.2.4. Fluidization Velocity
Fluidization velocity (fluidization is the processing technique employing a
suspension of a small solid particle in a vertically rising stream of fluid – usually
gas – so that fluid and solid come into intimate contact) affects the mixing of
particles within the reactor. Higher velocities increase the temperature of the fuel
bed and lead to the production of lower energy syngas.
28
3.2.5. Feed Properties
Both chemical reaction rates and diffusion rates are dependent on the
properties of the solid materials. The absolute value of the chemical reaction rate
can vary greatly depending on the reactivity of the material.
29
3.3. Block Diagram
Figure 3.1: Block Diagram
30
3.4. Process Flow Diagram
Figure 3.2: Process Flow Diagram
CHAPTER 4
MATERIAL BALANCE
4.1. Material Balance on Dryer
Total mass entering the dryer = F1 = 1000 kg/hr
Biomass : Coal = 30 : 70
Moisture content in Biomass = 15.3 kg/hr
Moisture content in coal = 333.2 kg/hr
Total moisture content = 348.5 kg/h
Dry Solid leaving dryer = F2 = 683.6 kg/hr
32
(There is only 5% moisture at the outlet of the dryer)
Where outlet of dryer contains
H2O = 32.1 Kg/hr
Coal + Biomass = 651.5 kg/hr
Mass In = Mass Out
Wet Solid In = Dry Solid out + Moisture removed
Moisture removed = 1000-683.6 = 316.4 Kg/hr
4.2. Material Balance on Mill
4.3. Material Balance on Gasifier
33
Table 4.1: Proximate analysis of Coal and Biomass (Dry Basis)
Coal [24] Biomass [10] Combined
Fixed Carbon 37.33% 18.03% 31.54%
Volatile matter 47.81% 71.9% 55.04%
Ash 14.86% 10.07% 13.42%
Total feed entering the gasifier = 683.6 kg/hr
Moisture content present in the feed = 32.1 kg/hr
Total Solid feed ( Coal + Biomass ) = 651.5 kg/hr
Assumptions
i. The gasifier operates under steady state conditions [21].
ii. Four reactions occur simultaneously in the gasifier including char
gasification, methane reforming, water gas shift and carbonation [22]
iii. Product gas consist of H2, CO, CO2 and CH4 [21]
iv. The reactions proceed isothermally and occur at constant volume [21]
v. Tar and ash formation in the process are negligible. As the calculation of
tar content will lead to an increasing amount of error for final product gas
composition [21]
4.3.1. Gasification Reactions
I. Char Gasification Reaction
By balanced chemical equation
Moles of C = 17.12 kmol/hr = 205.44 kg/hr
Moles of H2O = 17.12 kmol/hr = 308.16 kg/hr
Moles of CO = 17.12 kmol/hr = 479.36 kg/hr
Moles of H2 = 17.12 kmol/hr = 34.24 kg/hr
34
II. Steam Methane Reforming Reaction
By balanced chemical equation
Moles of CH4 = 22.41 kmol/hr = 358.56 kg/hr
Moles of H2O = 22.41 kmol/hr = 403.38 kg/hr
Moles of CO = 22.41 kmol/hr = 627.48 kg/hr
Moles of H2 = 67.23 kmol/hr = 134.46 kg/hr
III. Water Gas Shift Reaction
By balanced chemical equation
Moles of CO = 39.53 kmol/hr = 1106.84 kg/hr
Moles of H2O = 39.53 kmol/hr = 711.54 kg/hr
Moles of CO2 = 39.53 kmol/hr = 1732.32 kg/hr
Moles of H2 = 39.53 kmol/hr = 79.06 kg/hr
Amount of steam required = 1423.08 Kg/hr (S/C = 1)
For S:C = 2 (for high Hydrogen yield)
Amount of steam = 2846.16 Kg/hr (S/C = 2)
Table 4.2: Gasification Reaction
Name Reaction Conversion %
Char Gasification Reaction C(s) + H2O CO + H2 40 [23]
Methane Reforming Reaction CH4 + H2O CO + 3H2 60 [25]
Water Gas Shift Reaction CO + H2O CO2 + H2 70 [26]
Carbonation Reaction CaO + CO2 CaCO3 85 [20]
35
Table 4.3: Material Balance
Reactant Stream Products Stream Unreacted
C
(kmol/hr)
H2O
(kmol/hr)
CO
(kmol/hr)
H2 (kmol/hr) C (kmol/hr) H2O
(kmol/hr)
17.12 17.12 6.85 6.85 10.27 10.27
CH4
(kmol/hr)
H2O
(kmol/hr)
CO
(kmol/hr)
H2 (kmol/hr) CH4
(kmol/hr)
H2O
(kmol/hr)
22.41 22.41 13.45 40.35 8.96 8.96
CO
(kmol/hr)
H2O
(kmol/hr)
CO2
(kmol/hr)
H2 (kmol/hr) CO
(kmol/hr)
H2O
(kmol/hr)
20.3 20.3 14.21 14.21 6.09 6.09
CaO
(kmol/hr)
CO2
(kmol/hr)
CaCO3
(kmol/hr)
CaO
(kmol/hr)
CO2
(kmol/hr)
12.21 12.21 10.38 1.83 3.84
Total CaO fed = 683.6 kg/hr
Total CaCO3 at outlet = 1038 kg/hr
Unreacted CaO = 102.48 kg/hr
CO2 absorbed = 456.72 kg/hr
36
Table 4.4: Product Gas Composition of Gasifier
Component Molar flow rate (kmol /
hr)
Mass Flow Rate (kg/hr)
CH4 8.96 143.36
CO 6.09 170.52
CO2 3.84 168.96
H2 61.41 122.82
H2O 125.39 2256.55
Ash - 87.43
Unreacted C 10.27 123.24
Mass In = Mass Out
683.6 + 2846.16 + 683.6 = 1038 + 102.48 + 143.36 + 170.52 + 168.96 + 122.82
+ 2256.55 + 87.43 + 123.24
4213.36 = 4213.36
4.4. Material Balance on Cyclone Separator
Solid particles are coming in gas stream at the rate of 210.67 kg/hr. Particle Size
is reduced due to attrition. Assuming average particle size of particles in the Inlet
is 50 μm. From the graph the efficiency is 99%.
F6 = 210×0.99 = 208.56 kg/hr
F4 = F5 + F6
3072.88 = F5 + 208.56
Syn gas Out = F5 = 2864.32 Kg/hr
37
4.5. Material Balance on Scrubber
Table 4.5: Composition of Gas at Inlet of Scrubber
Component Molar flow rate (kmol
/ hr)
Mass Flow Rate
(kg/hr)
CH4 8.96 143.36
CO 6.09 170.52
CO2 3.84 168.96
H2 61.41 122.82
H2O 125.36 2256.55
Solids - 2.107
38
Water required = 33,400 kg/hr = 9.28 kg/s
Inlet temperature of water = 25 °C = 298 K
Outlet temperature of water = 89 °C = 362 K
Average Temperature of the Column = 61.4 °C = 334.4 K
Assuming only CO2 is absorbed in water, because solubility of other gases in
water is very less.
CO2 absorbed in water (At 61.4 °C) = 0.6 g CO2/Kg of water [27]
Total CO2 absorbed in water = 18 Kg/hr
Table 4.6: Outlet composition of Gases of Scrubber
Component Molar flow rate (kmol
/ hr)
Mass Flow Rate
(kg/hr)
CH4 8.96 143.36
CO 6.09 170.52
CO2 3.43 150.96
H2 61.41 122.82
H2O 0.49 8.82
Solids Removed = 2.107 kg/hr
Steam Removed = 2247.73 kg/hr
Total Outlet from the bottom = 33400+2247.73+18+2.107 = 35667.74 kg/hr
Total Gas Out from top = 596.48 kg/hr
Density of gas = 0.2953 kg/m3 [28]
Volumetric flow rate of the gas = 2020 m3/hr
39
4.6. Overall Material Balance
Figure 4.1: Overall Material Balance
CHAPTER 5
ENERGY BALANCE
5.1. Energy Balance on Dryer
5.1.1. Heat Inlet
Air inlet temperature = 385 K
Feed inlet temperature = 305 K
Air In
Air contains 0.01 kg water per kg of air
Q = mCp∆T
Heat Capacity of Air = 1.008 KJ/Kg K
Heat Capacity of Water Vapor = 1.869 KJ/Kg K
Q = 661.19 KW
41
Feed Inlet = Q = mCP∆T
Heat Capacity of Feed Solids = 1.104 KJ/Kg K
Heat Capacity of Water = 4.18 KJ/Kg K
Q = 19.36 KW
Total Heat In = Heat of Air at inlet + Heat of Feed at inlet
Total Heat In = 680.55 KW
5.1.2. Heat Outlet
Air outlet temperature = 327.5 K
Feed outlet temperature = 330 K
Feed Out
Q = mCp∆T
Q = 13.54 KW
Water evaporated = 0.008 Kg/s
Assuming evaporation takes place at 311 K
Latent heat of water vapour (λ) = 2410.8 KJ/Kg
Heat in water vapor = mCP∆T + mλ
= 228.834 KW
Heat in air = mCP∆T
= 321.74 KW
Heat losses = 116.43 KW
Total Heat Out = 564.12 KW
42
5.2. Energy Balance on Ball Mill
Energy In = Energy Out
Energy Out = 14.02 KW
5.3. Energy Balance on Gasifier
5.3.1. Heat Inlet
Solid In
Q = mCp∆T
Cp of feed = 1.1047 KJ/Kg. K
Cp of CaO = 0.7563 KJ/Kg. K
Q = 14.02+ 3.59 = 17.67 KW
Steam Inlet (At 523K and 2 atm)
Q = mCp∆T + mλ
Cp of Steam = 1.92 KJ/Kg. K
Cp of Water = 4.189 KJ/Kg. K
λ = 2200 KJ/Kg
At 2 atm, Steam boils at 393.57 K (120.57 ºC)
= (248.51+2200+505.18)*2846.16/3600
Q = 2335.19 KW
Total Heat In = Heat of Solid inlet + Heat Steam inlet
Total Heat In = 2352.86 KW
43
5.3.2. Heat involved in Chemical Reactions
I. Char Gasification Reaction
C + H2O → CO + H2
∆HºR = +131.5 KJ/mol
∆H = ∆HºR + ∆CP (T-TR)
∆CP = 2.276 KJ/Kmol .K
∆H = 133040.85 KJ/Kmol
II. Methane Reforming Reaction
CH4 + H2O → CO + 3H2
∆HºR = +206 KJ/mol
∆H = ∆HºR + ∆CP (T-TR)
∆CP = 12.39 KJ/Kmol .K
∆H = 214389.4 KJ/Kmol
III. Water Gas Shift Reaction
CO + H2O → CO2 + H2
∆HºR = -41 KJ/mol
∆H = ∆HºR + ∆CP (T-TR)
∆CP = 9.73 KJ/Kmol .K
∆H = -34414.14 KJ/Kmol
IV. Carbonation Reaction
CO2 + CaO → CaCO3
∆HºR = -178.3 KJ/mol K
∆H = ∆HºR + ∆CP (T-TR)
44
∆CP = 23.41 KJ/Kmol .K
∆H = -162451.43 KJ/Kmol
Table 4.7: Material Balance on Gasifier
Reactant Stream Products Stream Unreacted
C (kmol/hr)
H2O (kmol/hr)
CO (kmol/hr)
H2 (kmol/hr) C (kmol/hr) H2O (kmol/hr)
17.12 17.12 6.85 6.85 10.27 10.27
CH4
(kmol/hr) H2O
(kmol/hr) CO
(kmol/hr) H2 (kmol/hr) CH4 (kmol/hr) H2O (kmol/hr)
22.41 22.41 13.45 40.35 8.96 8.96
CO (kmol/hr)
H2O (kmol/hr)
CO2
(kmol/hr) H2 (kmol/hr) CO (kmol/hr) H2O (kmol/hr)
20.3 20.3 14.21 14.21 6.09 6.09
CaO (kmol/hr)
CO2
(kmol/hr) CaCO3
(kmol/hr) CaO (kmol/hr) CO2 (kmol/hr)
12.21 12.21 10.38 1.83 3.84
QR =∑ n∆H
=253.14 + 800.98 -135.84 – 468.40
= 449.88 KW
45
5.3.3. Heat Outlet
Total Heat Out = Qout = Qout (Gas) + Qout (Solids)
Heat Outlet by Gaseous stream = Qout (Gas) = m<Cp>∆T + mλ
λ = 2257 KJ/Kg (At 1 atm)
Mean Cp of Outlet Gaseous stream = 2.417 KJ/Kg K
Qout (Gas) = 2811.44 KW
Heat Outlet by Solids = Qout (Solids) = m<Cp>∆T
Mean Cp of Solids = 1.247 KJ/Kg K
Qout (Solids) = 267.45 KW
Total Heat Out = Qout = 3078.89 KW
5.3.4. Overall Energy Balance
Qin = Qout
QSteam + Qfeed + Qgeneration – Qconsumption + Qheater = Qout
Qheater = 1175.91 KW
Therefore Electric jacketed type heater containing electric wires is used to
maintain the temperature of the reactor.
5.4. Energy Balance on Cyclone Separator:
Energy In = 2811.44 KW
Energy Out = Qgas + Qsolid
Qsolid = m<Cp>∆T
Mean Cp of Solids = 1.252 KJ/Kg K
46
= 49.10 KW
Qgas = m<Cp>∆T
Mean Cp of Outlet Gaseous stream = 2.502 KJ/Kg K
Qgas = 2762.077 KW
Heat Inlet = Heat Outlet
2811.44 = 49.10 + 2762.077
5.5. Energy Balance on Scrubber:
Inlet temperature of gas = 677 ºC
Inlet temperature of water = 25 ºC
Table 5.1: Composition of Inlet Gas of Scrubber
Component Molar flow rate (kmol / hr) Mass Flow Rate (kg/hr)
CH4 8.96 143.36
CO 6.09 170.52
CO2 3.84 168.96
H2 61.41 122.82
H2O 125.39 2256.55
Solids - 2.107
Heat inlet by gas = Qgas= m<Cp>∆T = 2762.077 KW
Heat inlet by water = QWin = m<Cp>∆T
= 978.8 KW
Total Heat Inlet = 3740.88 KW
47
Table 5.2: Composition of Outlet Gases of Scrubber
Component Molar flow rate (kmol / hr) Mass Flow Rate (kg/hr)
CH4 8.96 143.36
CO 6.09 170.52
CO2 3.43 150.96
H2 61.41 122.82
H2O 0.49 8.82
Total 80.38 596.48
Heat out by gas = m<Cp>∆T = 23.44 KW
(Assuming that the scrubber cools the gas to 35 ºC)
Assuming no heat losses
Total Heat Out = 3740.88 KW
Heat Out from bottom = 3717.44 KW
Water required = Q = m<Cp>∆T
Assuming the outlet temperature of water = 95 °C
Water required = 33381.85 kg/hr
Actual flow rate of the water at top of column = 33,400 kg/hr
Outlet temperature of the water = m<Cp>∆T = 89 °C
Mass LHV of gas (KJ/Kg) = 3.963E+04 [28]
LHV of gas (KJ/Kgmol) = 2.959 E+05 [28]
48
5.6. Overall Energy Balance
Figure 5.1: Overall Energy balance
CHAPTER 6
EQUIPMENT DESIGN
6.1. Dryer Design
It includes the following steps:
Brief Introduction
Classification of dryers
Selection of dryer
Brief description of selected dryer
Design considerations
6.1.1. Brief Introduction
In general, drying a solid means the removal of relatively small amounts of
water or other liquid from the solid material to reduce the content of residual liquid
to an acceptably low value. Drying is usually the final step in a series of operations,
and the product from a dryer is often ready for final packaging. In a dryer, Water
or other liquids are removed from solids thermally by vaporization.
50
6.1.2. Classification of Dryers
Drying equipment may be classified in several ways. The two most useful
classifications are based on
1. The method of transferring heat to the wet solids.
2. The handling characteristics and physical properties of the wet material.
The first method of classification reveals differences in dryer operation and
design i.e.; it classifies dryers as direct or indirect, with subclasses of continuous
or batch wise operation. While the second method is most useful in the selection of
a dryer or group of dryers for a given drying problem.
6.1.3. Selection of dryer
From Classification No. 1
A. Direct Heat Dryers are selected, because
1. There is no limitation of reduced pressure in our drying problem.
2. Dusting problem is not serious.
3. The material to be handled is not heat sensitive around 100-150ºC.
B. Continuous Dryers are selected, because
1. For throughputs greater than 10,000 Kg/day, continuous dryers are used.
So Direct Heat Continuous Dryer is selected.
From Classification No. 2
C. Rotary Dryer is selected, because
1. Suitable for granular solids.
2. Used for High capacity/High production rate
3. Thermal Efficiency Range is 55—75%
4. Low operating Cost
5. Most Economical in Construction
51
6. Known as Workhorse of Process Industry
7. Low maintenance cost
So Direct Heat Rotary Dryer Is Selected.
6.1.4. Direct Heat Rotary Dryer
1. It consists of a cylinder, rotated upon suitable bearings and usually slightly
inclined to the horizontal.
2. Wet feed enters one end of the cylinder progress through it by virtue of
rotation and slope of the cylinder and discharge as finished product at the
other end.
3. In this dryer solids are directly exposed to hot gas usually air.
Figure 6.1: Direct Heat Rotary Dryer
A, dryer shell; B, shell-supporting rolls; C, drive gear; D, air-discharge hood; E,
discharge fan; F, feed chute; G ,lifting flights; H, product discharge; J air heater
52
6.1.5. Dryer Design [29]
Total mass entering the dryer = F1 = 1000 kg/hr
Biomass: Coal = 30:70
Moisture content in Biomass = 15.3 kg/hr
Moisture content in coal = 333.2 kg/hr
Total moisture content = 348.5 kg/hr
Solid feed rate leaving dryer = F2 = 684.07 kg/hr
(There is only 5% moisture at the outlet of the dryer)
Where outlet of dryer contains
H2O = 32.575 Kg/hr
Coal + Biomass = 651.5 kg/hr
Water removed =F3= 348.5 - 32.575
= 315.93 Kg/hr
Design Considerations
General design considerations for direct heat Rotary dryer are:
1. Solid Feed Rate and Moisture Content
2. Drum Diameter and Length
3. Drum Slope
4. Rotational Speed
5. Lifting Flights
6. Outlet Humidity
1. Solid Feed Rate and Moisture Content
Solid feed rate
F1=1000 kg/hr
Moisture Contents
Moisture contents=Mass of water/Mass of dry solid
Initial moisture content=348.5/651.5
Xa= 0.5349= 53.5 %
53
Final moisture content= 32.575/651.5
Xb= 0.05 = 5 %
Temperature Selection & Calculations
Temperature of air
1. Inlet air temperature (Ta )
1. Hot air is used for drying purpose and air inlet temperature varies from
100 to 150 oC
2. The higher the temperature of inlet air stream, the higher the
efficiency of the dryer in general. --- (a)
3. So take inlet air temperature= Ta = 112 ºC = 385K
Humidity of Inlet air (H1)
Inlet humidity of air is found with the help of dry bulb temperature
(Room temp) of air and wet bulb temperature of entering air from
Humidity chart.
i.e.; At Tdb = 32oC = 305K
Twb = 21oC = 294K
Humidity of air=H1= 0.01 kg of water/kg of dry air
Now this air is heated from 32oC to 112oC (dry bulb temperature of heated
air) and the wet bulb temperature of heated air is found with the help of dry bulb
& humidity of heated air entering the dryer from humidity chart.
i.e.; At Tdb = 385K
H1= 0.01 kg of water/kg of dry air
Twb = 311K
Outlet air temperature (Tb)
The proper outlet gas temperature is a matter of economics,
54
It is estimated from
---------------- (1)
Where
Nt =No. of transfer units (1.5—2.5 for rotary dryers)
Ta = 385K
TWb = 311K
Take Nt =1.5, putting the values in Eq. (1)
Air outlet temperature =Tb = 327.5K
Temperature of Feed
a) Inlet feed temperature (Tsa)
Feed enters the dryer at Tsa = 305K
b) Outlet Feed temperature (Tsb)
Feed leaves the dryer at Tsb = 330K
Temperature pattern in continuous countercurrent adiabatic dryer:
Feed inlet temperature should be less than gas outlet temperature and it should
also less than wet bulb temperature of air, therefore selected feed inlet
temperature is correct.
Outlet temperature of feed is slightly greater than gas outlet temperature,
therefore selected solid outlet temperature is correct.
Table 6.1: Temperature of Inlet and Outlet Streams
Material
Inlet
Temperature K
Outlet
Temperature K
Feed 305 330
Air 385 327.5
Wbb
Wba
tTT
TTLnN
55
2. Drum Diameter
To calculate diameter of the rotary dryer, following steps must be performed.
i. Calculate heat duty on dryer (qt)
ii. Calculate mass flow rate of air entering (mg )
iii. Choose the air mass velocity (G)
iv. Find the dryer area (A)
i. Calculation of heat load on dryer (qt)
Heat must be supplied to a dryer to accomplish the following
A. Heat the feed (solid & liquid) to the vaporization (wet bulb) temperature
B. Vaporize the liquid.
C. Heat the solids to their final temperature.
D. Heat the vapor to its final temperature.
In general, rate of heat transfer per unit mass can be calculated as follows
----------------- (2)
Where
qt =Rate of heat transfer
ms =1000 kg/hr
Tsa = 305K
Tsb =330K
Ta =385K
Tb =327.5K
TWb = 311K
Xa = 0.5349 = 53.5%
Xb= 0.05= 5 %
)(
)()()(
)()(
wbsblb
wbbvbaba
sawbLasasbss
TTCpX
TTCpXXXX
TTCpXTTCpm
qt
56
From steam tables, latent heat of vaporization (λ) =2410.8 KJ/kg (At 311K)
Heat capacity of feed stock = 0.3 Cp bagasse + 0.7 Cp coal
Heat capacity of bagasse [30] = 0.46KJ/kg K
Heat capacity of coal [31] = 1.381KJ/kg K
Heat capacity of solid feed stock = Cps = 1.104KJ/kg K
Heat capacity of water [32] = CpL = 4.178 KJ/kg K
[At T= (311+305)/2]
Heat capacity of water vapors [33] = Cpv = 1.869 KJ/kg K
[At T= (311+327.5)/2]
Putting these values in Eq. (2),
qt /ms = 1230.41 KJ/kg
qt = 1228.41*0.2777 KJ/s =341.38KW
ii. Calculation for air flow rate (mg)
Mass flow rate of air entering is found from given formula
------ (3)
Where
Csb=Humid Heat = 1.005 + 1.82H KJ/kg K = 1.0232 KJ/kg K
H1=Inlet humidity
Putting the values in Eq. (3),
mg = 5.74 Kg/s
iii. Air mass velocity (G)
The allowable mass velocity of air in direct contact rotary dryer ranges from 2000
to 25000 kg/m2 hr.
Take G = 12000 kg/m2 hr = 3.4 kg/m2 s
)1)(( 1HTTC
qg
absb
tm
57
iv. Area of dryer (A)
It can be found as
Area of dryer = Air mass flow rate = 1.724 m2
Air mass velocity
v. Drum Diameter
Diameter is calculated by equation
D=1.483 m
vi. Drum Length
To calculate Length we have to go through the following procedure
i. Calculate the Volumetric heat transfer coefficient (Ua)
ii. Calculate log mean of drying gas wet bulb depression at inlet & exit end of
dryer (ΔTm)
iii. Calculate the volume of dryer (V)
vii. Volumetric heat transfer coefficient (Ua)
Volumetric heat transfer Coefficient is found from given formula
Ua =0.756 KW/m3 K
5.0*4
AD
D
GaU
67.0^5.0
58
viii. Calculation for log mean of drying gas wet bulb depression at inlet
& exit end of dryer (ΔTm)
Log mean depression Of Wet bulb (ΔTm) is found from the following formula
ΔTm= 38.32K
ix. Volume of dryer (V)
Volume of dryer is found from the given formula
V=11.78 m3
x. Length of dryer
Length of dryer is calculated as
L = volume of dryer / area of dryer
L = 11.78 / 1.72 = 6.83 m
L/D ratio
L/D= 6.83/1.72
= 4.6
As L/D ratio for rotary dryer varies from 1 --- 10, therefore the above values can
be accepted.
)]/()[(
)()(
wbawbb
wbawbb
TTTTLn
TTTTTm
ma
t
TU
QV
*
59
2. Slope of drum
Slope of drum should be kept from 0 to 8 cm/m or 0 to 5°, more the slope of the
drying drum more will be forward driving force but product abrasion will also
increase. The slope is taken as 5 cm/m=3°
3. Rotational Speed
Assume the peripheral speed of rotation to be 06 m/min
Revolutions per minute = Peripheral Speed/Diameter
RPM = 6/1.48 = 4.04
= 4.04
The revolution of a drier varies between 2– 5
Therefore the above value can be accepted.
4. Lifting Flights
No. of Flights
The standard no. of flights is 2 to 4 times the diameter
No. of Flights=3*D
=4.5 ≈ 5
Radial height of the flight
The flight heights in direct rotary dryer will range from 1/12 to 1/8th of dryer
diameter.
Flight heights= 1/8*Diameter
=0.185 m
60
Shape of Flights
The shape of the flights depends upon the handling
characteristics of the solids. For free flowing materials, a radial
flight with a 900 lip is employed. So the flights are of radial
type.
5. Outlet Humidity
The air outlet humidity is found from given formula
Where
H2= Outlet Humidity
H1= Inlet Humidity
=0.01 kg of water/ kg of dry air
ms = 1000 kg/hr
mg =20700 kg/hr
Xa = 0.535
Xb= 0.05
Putting the values in above equation, we get
H2=0.0334 Kg of water/Kg of dry air
g
bas
m
XXmHH
)(12
61
6.1.6. Dryer Specification Sheet
6.1.7. Dryer Design Data
Equipment Dryer
Function To reduce the Water contents
Operation Continuous
Type Direct heated Rotary dryer
Flow rate of Solid entering 1000 Kg/hr
Inlet moisture content 53.5%
Outlet moisture content 5%
Mass flow rate of Air 20700 kg/hr
Total heat requirement 341.38 KW
Air mass velocity 3.33 Kg/s-m2
Area of dryer 1.723 m2
Diameter of dryer 1.48 m
Length of dryer 6.83 m
Volume of dryer 11.78 m3
L/D ratio 4.6
62
6.2. Gasifier Design
6.2.1. Fluidized Bed Reactor [22]
The designing of fluidized bed reactor consists of following steps.
1. Calculation of minimum fluidization velocity (Umf)
2. Calculation of Steam flow rate (Qo)
3. Calculation of Reactor Diameter, height and volume
4. Distributor plate design
6.2.2. Hydro Dynamics
Figure 6.2: Hydrodynamics Calculations [22]
6.2.3. Minimum Fluidization Velocity
Minimum fluidization velocity is calculated from the hydrodynamic study
based on physical properties of bed particles and fluidizing gas particles (Table 6.1
and Table 6.2). Bed particles consists of coal, sugarcane bagasse and CaO, whereas
CaO is used as CO2 sorbent in the bed. CaO is the heaviest among the bed particles
and considered as single bed particle in reactor dimensions evaluation. Moreover,
CaO mean particle diameter is initially taken as a single diameter in the gasifier.
63
Table 6.1: Bed Material Properties [22]
Bed material CaO
Mean particle diameter (dp) 2 mm
Particle density (Ƿs) 2551kg/m3
Bulk density (Ƿ) 1089kg/m3
Bed voidage (Emf) 0.57
Sphericity (Øs) 0.43
Table 6.2: Fluidizing Gas Properties (at 950K and 1 atm)
Fluidizing Gas Steam
Density (Ƿf) 0.2312kg/m3
Viscosity (µ) 0.0000356kg/m.s
At inlet the steam pressure must be 2 atm to maintain 1 atm pressure inside the
reactor (Assume that temperature of steam at the inlet is equal to the temperature
of the steam inside the reactor)
Table 6.3: Fluidizing Gas Properties (at 523K and 2 atm)
Fluidizing Gas Steam
Density (Ƿf) 0.8449kg/m3
Viscosity (µ) 0.00001818kg/m.s
Minimum fluidization velocity can be calculated from the following equation.
------------- (1)
Whereas Reynolds number at minimum fluidization is calculated from Ergun’s
(1952) equation.
------------- (2)
Where Ar is Archimedes number and can be calculated as:
fp
mf
mfd
U
Re
2
332Re
75.1Re
1150mf
mfs
mf
mfs
mfAr
64
------------- (3)
Where
Remf = Reynold’s number at minimum fluidization
g = Gravitational acceleration
By putting the values in Eq. (3)
Ar = 511093
By putting the values in Eq. (2)
Ar = 150(1 – Emf)Remf/Øs2Emf
3 + 1.75 Remf2/ØsEmf
3
511093 = [150(1 – 0.57)Remf / (0.43)2(0.57)3] + [1.75Remf2 / (0.43)(0.57)3]
511093 = 21.976 Remf2 + 1883.64 Remf
0 = Remf2 + 85.714 Remf – 23256.87
By using Quadratic formula
Remf = (-85.714+316.82)/2 Remf =(-85.714-316.82)/2
Remf = 115.55 Remf = -201.27 (not possible)
Using the value of Reynold’s number in Eq. (1)
Umf = 1.24 m/s
2
3
gdAr
fsfp
65
6.2.4. Diameter of Reactor
Reaction based steam Calculations
C(s) + H2O CO + H2 ------------- (4)
(Char Gasification reaction)
CH4 + H2O CO + 3 H2 ------------- (5)
(Steam Methane Reforming)
CO + H2O CO2 + H2 ------------- (6)
(Water Gas shift reaction)
Table 6.4: Proximate analysis of Coal and Biomass (Dry Basis)
Coal [24] Biomass [10] Combined
Fixed Carbon 37.33% 18.03% 31.54%
Volatile matter 47.81% 71.9% 55.04%
Ash 14.86% 10.07% 13.42%
Total feed entering the gasifier = 683.6 kg/hr
Moisture content present in the feed = 32.1 kg/hr
Total Solid feed ( Coal + Biomass ) = 651.5 kg/hr
In 651 kg of feed
C=205.5 kg = 17.12 kmol
CH4 = 358.6 kg = 22.41 kmol (For steam calculations, assume that all VM = CH4)
Ash = 87.43 kg
66
Char Gasification Reaction
By balanced chemical equation
Moles of C = 17.12 kmol/hr = 205.44 kg/hr
Moles of H2O = 17.12 kmol/hr = 308.16 kg/hr
Moles of CO = 17.12 kmol/hr = 479.36 kg/hr
Moles of H2 = 17.12 kmol/hr = 34.24 kg/hr
Steam Methane Reforming Reaction
By balanced chemical equation
Moles of CH4 = 22.41 kmol/hr = 358.56 kg/hr
Moles of H2O = 22.41 kmol/hr = 403.38 kg/hr
Moles of CO = 22.41 kmol/hr = 627.48 kg/hr
Moles of H2 = 67.23 kmol/hr = 134.46 kg/hr
Water Gas Shift Reaction
By balanced chemical equation
Moles of CO = 39.53 kmol/hr = 1106.84 kg/hr
Moles of H2O = 39.53 kmol/hr = 711.54 kg/hr
Moles of CO2 = 39.53 kmol/hr = 1732.32 kg/hr
Moles of H2 = 39.53 kmol/hr = 79.06 kg/hr
Amount of steam required = 1423.08 Kg/hr (S/C = 1)
For S:C = 2
Amount of steam = 2846.16 Kg/hr
Density of steam = 0.2312 Kg/hr
67
Volumetric flow rate of steam = Qo = 12310.38 m3/hr ≈ 12310 m3/hr
= 3.42 m3/s
Uo = 4 m/s (Uo = 3—5 Umf)
21
02Re
oU
QIDactor
Reactor ID = 1.043 m ≈ 1.04 m
6.2.5. Height of Reactor
Height of the reactor is calculated as
Total Height = Bed height + TDH
For stable operation
Bed Height = Bed Diameter
Bed Height = 1.04m
Total disengaging Height (TDH)
TDH = 4.47 [Dbm]0.5
Dbm = 0.652 [At (U – Umf)]2/5
Uo = 4 m/s (Uo = 3—5 Umf)
𝐴𝑡 = 𝜋𝐼𝐷2
4
= 0.85 𝑚2
𝐷𝑏𝑚 = 0.652 [ 0.85 (4 − 1.24)]2/5
𝐷𝑏𝑚 = 0.917 𝑚
𝑇𝐷𝐻 = 4.47 [𝐷𝑏𝑚] 0.5
TDH = 4.281 m ≈ 4.28 m
68
Total Height = 1.04 + 4.28 = 5.32 m ≈ 5.3 m
6.2.6. Volume of Reactor
Volume = π*r2*L
= 4.5 m3
6.2.7. Distributor Plate Design [34]
𝑅𝑐 = 𝐷𝑖𝑠𝑡𝑟𝑖𝑏𝑢𝑡𝑜𝑟𝑃𝑟𝑒𝑠𝑠𝑢𝑟𝑒𝐷𝑟𝑜𝑝
𝐵𝑒𝑑𝑃𝑟𝑒𝑠𝑠𝑢𝑟𝑒𝐷𝑟𝑜𝑝=0.01+0.2[1 − 𝑒−0.5𝐷/𝑍]
R c =0.01+0.2[1 − 𝑒−0.5×0.9/0.9]
R c=0.089
ΔPb =H (1-ε) (ρs-ρf)
ΔPb = 5.3 (1-0.57)(2551-0.8449)
ΔPb= 5811.80 kg
m2 = 56955.67 Pa
ΔPd= Rc ΔPb = 0.089 x 5811.80 = = 517.25 kg
m2 = 5069.70 Pa
69
6.2.8. Number of Orifices in the Distributor Plate [35]
Vessel Reynold’s number can be calculated by the following formula.
Re = ρfxdxUo
μ
= 0.2312 * 1.04 * 4 / 0.0000356 = 27016.63
Re 100 300 500 1000 2000 >3000
CdOR 0.68 0.70 0.68 0.64 0.61 0.60
From the table CdOR = 0.6
UOR=Cdor(2×∆𝑝𝑑
𝜌𝑓)0.5
= 0.6 (2×5069.7
0.8449)0.5
UOR = 65.73 m/s
𝑈𝑜 = 𝐴𝑜𝑟 ∗ 𝑈𝑜𝑟 ∗ 𝑁𝑜𝑟
Whereas
𝑁𝑜𝑟 =𝑈𝑜
𝐴𝑜𝑟 ∗ 𝑈𝑜𝑟
𝑁𝑜𝑟 =4
𝐴𝑜𝑟 ∗ 65.73
NOR = 0.06/ AOR
Dia of orifice = 2 mm
Area of orifice = 3.1416 E-6
NOR = 19108
70
6.1.9. Gasifier Specification Sheet
Equipment Gasifier
Function To Convert Solid fuel into Gaseous Product
Operation Continuous
Type Fluidized Bed Gasifier
6.2.10. Gasifier Design Data
Minimum Fluidizing Velocity 1.24 m/s
Volumetric flow rate of steam 3.42 m3/s
Reactor ID 1.04 m
Bed Height 1.04 m
TDH 4.28 m
Total Height 5.3 m
Volume of Reactor 4.5 m3
Diameter of Orifice 2 mm
Area of Orifice 3.1416 E-6
NOR 19108
71
6.3. Cyclone Separator Design [36]
6.3.1. General design procedure
1. Select either the high-efficiency or high-throughput design, depending on
the performance required.
2. Estimate the number of cyclones needed in parallel.
3. Calculate the cyclone diameter for an inlet velocity of 9-27 m/s. Scale the
other cyclone dimensions from Figure.
4. Calculate the scale-up factor.
5. Calculate the cyclone performance and overall efficiency (recovery of
solids). If unsatisfactory try a smaller diameter.
6. Calculate the cyclone pressure drop and, if required, select a suitable blower
6.3.2. Number of Cyclones
Flow rate of the gas entering = 3072.88 kg/hr
Density = 0.1922 [28]
Volumetric flow rate = mass/density
= 15980 m3/hr
= 4.44 m3/s (High throughput design is selected)
Standard velocity taken in cyclone is = u = 9-27 m/s
Figure 5.3: High through-put Cyclone
72
6.3.3. Inlet Duct area
Area= Volumetric Flow rate / Velocity
Area of inlet duct at 20 m/s =4.44/20
Area = 0.2219 m2
DC:
From the Figure
Duct area= 0.75Dc × 0.375 Dc
0.2219 = 0.28125 ×𝐷c2
Dc= 0.888 m
𝑆𝑡𝑎𝑛𝑑𝑎𝑟𝑑 Dc = 0.203 m
0.888/0.203 = 4.37
4 cyclones should be used in parallel.
Flow rate per cyclone = 3995 m3/h
6.3.4. Dimensions of Cyclone
Total Height:
0.888/2=0.444m =Dc
From the figure.
Total height = 1.5×𝐷c + 2.5×𝐷c
= 4×0.444= 1.776m
Outlet Duct Area:
From the figure
Do = 0.75×𝐷c
= 0.75×0.444
= 0.333 m
Ao = 𝜋×(0.333/2)2
= 0.087 m2
73
Dust Exit Diameter:
From the Figure
Dust exit diameter = 0.375 × Dc
= 0.375 ×0.444
= 0.166 m
6.3.5. Calculation of Scaling Factor
d2/d1 = [(𝒅𝒄𝟐/𝒅𝒄𝟏)𝟑×(𝑸𝟏/𝑸𝟐)×(Δ𝝆𝟏/ Δ𝝆𝟐 )×(𝝁𝟐/𝝁𝟏)]𝟎.𝟓
dc2 = Diameter of cyclone=0.444 m
dc1 = Diameter of standard cyclone=0.203 m
Q1 = Standard flow rate= 669 m3/h
Q2 = Volumetric flow rate=3995 m3/h
Δ𝝆𝟏 = standard solid-fluid =2000 kg/m3
Δ𝝆2 = particle density=1642 kg/m3
𝝁1 =standard viscosity =0.018 mNs/m2 (cp)
𝝁𝟐 =gas viscosity = 0.02638 cp (mNs/m2)
By putting these values.
d2/d1 = [(0.444/0.203)𝟑×(669/3995)×(2000/ 1642 )×(0.02638/0.018)]𝟎.𝟓
d2/d1 = 1.75
Therefore it has 99% efficiency when the particle size at the outlet is 0.05 mm
74
6.3.6. Pressure Drop Calculation
ΔP = (𝛒𝐟
𝟐𝟎𝟑) {𝐮𝟏𝟐 [𝟏 + 𝟐𝛗𝟐 ((
𝟐𝐫𝐭
𝐫𝐞) − 𝟏)] + (𝟐𝐮𝟐𝟐)}
ΔP=Cyclone pressure drop, millibars
𝝆f = Gas Density = 0.1790 kg/m3
u1 = 20 m/s
u2 = Volumetric Flow rate / Area of exit pipe = 1.11/0.087 = 12.75 m/s
rt = Radius of the circle to which the center line of the inlet is tangential
= [.444-(.166/2)]/2
= 0.18036 m
= 180.36 mm
re = Radius of the exit pipe = .16648m = 166.48 mm
Ψ = (fc × AS)/Ai
fc=0.005 for gases
Surface area of cyclone = As = Dc×ℎ×𝜋
=0.444×1.776×3.14
=2.47 m2
Ai=0.75 Dc* 0.375 Dc
= 0.051 m2
Ψ = 0.24
From the graph
φ = 1.1
ΔP = (𝟎.𝟏𝟕𝟗
𝟐𝟎𝟑) {(𝟐𝟎)𝟐[𝟏 + 𝟐(𝟏. 𝟏)𝟐(𝟐. 𝟏𝟔𝟔 − 𝟏)] + (𝟐(𝟏𝟐. 𝟕𝟓)𝟐)}
ΔP = 1.63 millibar = 16.72 mm H2O
Pressure drop is reasonable.
75
6.3.7. Cyclone Separator Specification Sheet
Equipment Cyclone Separator
Function To separate undesired particles
Operation Continuous
Type High Throughput Design Separator
6.3.8. Cyclone Separator Design Data
Flow rate of Entering gas 3072.88 kg/hr
Inlet Velocity 20 m/s
Inlet Duct Area 0.222 m2
Diameter of cyclone 0.444 m
Number of Cyclones 4
Flow Rate of each Cyclone 3995 m3/h
Total Height 1.77 m
Outlet duct Area 0.087 m2
Dust Exit Diameter 0.166 m
Gas Density 0.1790 kg/m3
Radius of Exit Pipe, re 166.48 mm
Pressure Drop 1.63 millibar
76
6.4. Design of Scrubber
6.4.1. Designing Steps
1. Selection of column.
2. Selection of packing and material
3. Calculating the size of packing
4. Determining the no. of transfer units (NOG)
5. Calculating the diameter of column
6. Determining the height of packing
7. Determining the height of the column
8. Determining the Column Pressure drop
Table 6.5: Scrubber Material Balance
Feed composition
Components Molar flow rate
(kmol / hr)
Mass Flow Rate
(kg/hr)
Mole Fraction
CH4 8.96 143.36 0.0436
CO 6.09 170.52 0.0296
CO2 3.84 168.96 0.0187
H2 61.41 122.82 0.2986
H2O 125.36 2256.55 0.6095
Solids - 2.107
Total 205.66 2864.317 1
77
Table 6.6: Scrubber Top Composition
Top Composition
Component Molar flow rate
(kmol / hr)
Mass Flow Rate
(kg/hr)
Mole fraction
CH4 8.96 143.36 0.1115
CO 6.09 170.52 0.0757
CO2 3.43 150.96 0.0427
H2 61.41 122.82 0.7640
H2O 0.49 8.82 0.0061
Total 80.38 596.48 1.00
Gm(y1 – y2) = Lm(x1 – x2) [37]
Gm=flow rate of gas entering (Kg/hr)
Lm = flow rate of solvent entering (Kg/ hr)
Y1=Mole fraction of Steam in entering stream of gas
Y2= Mole fraction of Steam in leaving stream of gas
X1= Mole fraction of Steam in leaving solvent stream
X2= Mole fraction of Steam in entering solvent stream
2864.317(Y-0.0061) = 33400(X-0)
Y= 11.66X +0.0061 ……………… (1)
Equation (1) is the operating line equations.
Equation for Equilibrium Curve:
78
Y1= Mole fraction of Steam in entering stream of gas = 0.61
Y2= Mole fraction of Steam in leaving stream of gas = 0.0061
At Y1= 0.61, using Equation (1) X1= 0.052
Using the figure the equation for equilibrium curve is;
ye= 7.5x ……………… (2)
At X1= 0.052 using equation (2), ye1= 0.39
m Gm′
Lm′ =
ye1
y1 −y2 =
0.39
0.61−0.0061
m Gm′
Lm′ = 0.646
y1
y2 =
0.61
0.0061 = 100
From figure 12.23 [39]
At y1
y2 = 100 and m
Gm′
Lm′ = 0.646
NOG= 10.1
6.4.2. Calculation of Diameter of Column [36]
Flow rate of entering gases = G = 2864.32 Kg/hr = 2864.32
3600 = 0.796 Kg/sec
Flow rate of entering solvent = L = 33400 Kg/hr = 33400
3600 = 9.28 Kg/sec
Temperature of entering gas = Tg = 677 °C = 950K
Temperature of entering Solvent = TL= 25 °C = 298K
Average column temperature = Tavg = TDG + 𝐶𝑝𝑔(𝑇𝑤−𝑇𝑔)+(1−ѡ)∆𝐻𝑎
𝐶𝑝𝑔+𝐶𝑝𝑤
Tavg = 61.4 ºC
Operating pressure of Scrubber = P = 1 atm
79
Average molecular weight of entering gases = 13.96 Kg
Density of gas mixture = ρg = 1.033 Kg/m3 [28]
Density of liquid solvent = ρL = 0.9793 g/cm3 = 979.3 Kg/m3
Viscosity of liquid solvent = µL =0.4534 cp = 4.534 × 10−4 Ns/m2 [28]
Viscosity of Gaseous mixture = µg = 0.0174 cp = 1.74 × 10−5 Ns/m2 [28]
Now
Abscissa of fig 11.44 [36]
Lw
Gw√
ρg
ρL =
9.28
0.796√
1.033
979.3 = 0.378
For pressure drop 40 mm of H2O /m of packing
From fig 11.44 [36]
K4 = 0.7 at 40mm H2O
And
K4 = 1.5 at flooding
% Flooding = √0.7
1.5 × 100 = 68.31 % (Satisfactory)
80
Packing Specifications
Packing Type = Ceramic Intallox Saddles
Packing Size =0 38 mm
Bulk density = 625 kg/m3
Surface area = a = 194m2/m3
Packing factor = Fp = 170 m-1
From equation 11.118
Vm∗ = [
K4 ρv(ρL − ρv)
13.1 FP(μLρL
)0.1 ]
0.5
Vm∗ = [
0.7×1.033(979.3− 1.033)
13.1×170(4.534×10−4
979.3)
0.1]
0.5
Vm∗ = 1.336 Kg/m2.s
As
Column Area required = A = G
Vm∗ =
0.796
1.336 m2 = .595 m2
Diameter of column =D =√ 4A
π
Diameter of column= 0.9 m
Actual Area of column = 0.636 m2
%age flooding (at 0.9m dia) = 73.02 %
81
6.4.2. Calculation of Height of Transfer Units (Onda’s Method)
Equation for calculation of effective interfacial area is given as.
Where
aw = effective interfacial area of packing per unit volume m2/m3
Lw = liquid mass velocity kg/m2s = 14.59 kg/m2s
a = actual area of packing per unit volume m2/m3 = 194 m2/m3
σc = critical surface tension for metal packing material = 61 x 10-3 N/m
σL = liquid surface tension N/m = 65.71 dyne/cm = 65.71 x 10-3 N/m
µL= 4.534 × 10−4 Ns/m2
ρL = 979.3 Kg /m3
By putting these values
aw = 142.71 m2/m3
6.4.3. Calculation of Liquid Film Mass Transfer Coefficient
KL = liquid film coefficient m/s
dp = packing size =38x 10-3 m
DL = diffusivity of liquid = 6.4 x 10-9 m2/s
Then, by substituting the values,
KL = 8.16 x 10-4 m/s
2.02
05.0
2
21.075.0
45.1exp1a
L
g
aL
a
L
a
a
LL
w
L
w
L
w
l
cw
4.02
1
3
2
3
1
0051.0 p
LL
L
Lw
w
L
LL ad
Da
L
gK
82
6.4.4. Calculation of Gas Film Mass Transfer Coefficient
Where KG = Gas film coefficient, kmol/m2s.bar
VW = Gas mass velocity = 1.25 Kg/m2.s
K5= 5.23 (For packing size above 15mm, Coulson & Richardson)
Dv =Diffusivity of gas = 5.43 x 10-5 m2/s
Then, by substituting the values,
KG = 1.85 x 10-3 kmol/m2s.bar
6.4.5. Calculation of Gas-film Transfer Unit Height
Where,
HG = Gas-film transfer unit height
Gm = 1.25/13.96 = 0.089 Kmol/m2.sec
Then,
HG = 0.089
1.85 x 10−3×142.71×1.013
HG = 0.333 m
23
17.0
5
p
gg
g
g
w
g
gGad
Da
VK
aD
RTK
PaK
GH
WG
mG
83
6.4.6. Calculation of Liquid-film Transfer Unit Height
Where
HL= Liquid-film transfer unit height
Lm= 14.59/18 = 0.81 Kmol/m2.s
Ct = Concentration of solvent = 979.3/18 = 54.4 Kmol/m3
Then,
HL = 0.81
8.16x 10−4×142.71×54.4
HL = 0.128 m
6.4.7. Calculation of Height of Transfer Units
As,
HG = 0.333 m
HL = 0.128 m
So,
Height of transfer units=HOG = 0.333 + 0.646 × 0.128
HOG = 0.42 m
6.4.8. Calculation of Height of Tower
Total height of packing = Z = NOG × HOG
Z = 10.1 × 0.42 = 4.242 m = 4.25 m
Allowances for liquid distribution = 1m
Allowances for liquid Re-distribution =1m
Total height of tower = 4.25 + 1 + 1
LHmL
mmGGHoGH
tWL
mL
CaK
LH
84
Total height of tower = Zt = 6.25
6.4.9. Pressure Drop Calculations
Pressure drop in a gas absorber is a function of superficial gas velocity and
properties of the packing used. The Scrubber may be designed for a specific
pressure drop or pressure drop may be estimated using Leva’s correlation.
∆P = c(10𝑗∗𝐿) ∗ ((𝑓 ∗ 𝐺𝑣)2/ρv)
ΔP: Pressure drop in inches water/foot packing
Gv: Superficial gas velocity = 0.256 lb/hr-s2
𝜌v: Gas density in lb/ft3 = 0.064 lb/ft3
L: Liquid superficial velocity = 2.99 lb/hr-s2
𝜌L: Liquid density in lb/ft3 = 61.13 lb/ft3
c = 0.14
j = 0.14
∆P = 0.2 inch of H2O/ft of packing
85
6.4.10. Scrubber Specification Sheet
Equipment Scrubber
Function Removal of Steam
Operation Continuous
Type Packed Column
Packing Type Ceramic Intallox Saddles
6.4.11. Scrubber Design Data
Flow rate of Entering gas 2864.32 kg/hr
NOG 10.1
Diameter of Column 0.9 m
Percentage Flooding 73.02
Effective Packing Interfacial Area 142.71 m2/m3
Liquid Film Mass Transfer Coefficient 8.16 E-4 m/s
Gas Film Mass Transfer Coefficient 1.85 E-3 kmol/m2 s.bar
HG 0.333 m
HL 0.128 m
HOG 0.42 m
Height of Packing 4.25 m
Height of Column 6.25 m
∆P 0.2 inch H2O/ft of packing
86
CHAPTER 7
INSTRUMENTATION AND PROCESS CONTROL
7.1. Instrumentation
Measurement is a fundamental requisite to process control. Either the
control can be affected automatically, semi-automatically or manually. The quality
of control obtainable also bears a relationship to the accuracy, re-product ability
and reliability of the measurement methods, which are employed. Therefore,
selection of the most effective means of measurements is an important first step in
the design and formulation of any process control system.
7.2. The Concept of Measurement in Automation Application
Measurement is defined as the extraction from physical and chemical
systems or processes of signals, which represent parameters or variable. The
performance of an automation system can never surpass that the associated
measuring devices. A basic example is a human being. The output of a measuring
instrument that has its output compared to an arbitrarily chosen reference of
suitable magnitudes which is normally assumed to be unvarying. “Transducer” or
“Sensor” is a general term for a sensing device.
87
7.3. Measurement
Instrumentation can be used to measure certain field parameters (physical
values). These measured values include:
1. Pressure
2. Flow
3. Temperature
4. Level
5. Density
6. Viscosity
7. Radiation
8. Frequency
9. Current
10. Voltage
7.4. Process
A process is broadly defined as an operation that uses resources to transform
inputs into outputs. It is the resource that provides the energy into the process for
the transformation to occur.
7.4.1. Process Control
Process control is the act of controlling a final control element to change the
manipulated variable to maintain the process variable at a desired Set Point.
88
7.4.2. Objectives of Process Control
Effective process control objective is to maintain safe operations, quality
products, and business viability.
I. Safety
The primary purpose of a Process Control system is safety: personnel safety,
environmental safety and equipment safety. The safety of plant personnel and the
community is the highest priority in any operation. An example of safety in a
common heat exchanger process is the installation of a pressure relief valve in the
steam supply.
II. Quality
In addition to safety, process control systems are central to maintaining
product quality. In blending and batching operations, control systems maintain the
proper ratio of ingredients to deliver a consistent product. They tightly regulate
temperatures to deliver consistent solids in cooking systems. Without this type of
control, products would vary and undermine quality.
III. Business Viability
When safety and quality concerns are met, process control objectives can
be focused on profit. All processes experience variations and product quality
demands that we operate within constraints. A batch system may require +- 0.5%
tolerance on each ingredient addition to maintain quality. A cook system may
require +- 0.5 degrees on the exit temperature to maintain quality. Profits will be
maximized the closer the process is operated to these constraints. The real
challenge in process control is to do so safely without compromising product
quality.
89
7.5. Basic Elements of Process Control
Controlling a process requires knowledge of four basic elements, the
process itself, the sensor that measures the process value, the final control element
that changes the manipulated variable, and the controller.
I. Process
It represents the material equipment together with physical or chemical
operation that occurs.
II. Sensors
Sensors measure the value of the process output that we wish to effect. This
measurement is called the Process Variable or PV. Typical Process Variables that
we measure are temperature, pressure, mass, flow and level.
III. Final Control Element
A Final Control Element is the physical device that receives commands
from the controller to manipulate the resource. Typical Final Control Elements
used in these processes are valves and pumps.
IV. The Controller
This is the hardware element that has “intelligence”. It receives the
information from the measuring device and decides what action must be carried
out. The older controllers were of limited intelligence, could perform very limited
and simple operations and could implement very simple control laws. The use of
digital computers in this field has increased the use of complicated control laws.
90
7.6. Basics of Process Control
1. Open Loop Control
2. Closed Loop Control
I. Open Loop Control
In open loop control we are not concerned that a particular Set Point be
maintained, the controller output is fixed at a value until it is changed by an
operator. Many processes are stable in an open loop control mode and will maintain
the process variable at a value in the absence of a disturbance.
II. Closed Loop Control
In closed loop control the controller output is determined by difference
between the process variable and the Set Point. Closed loop control is also called
feedback or regulatory control.
7.7. Selection of Controller
Actually in industry, only P, PI and PID control modes are the usual
practice. The selection of most appropriate type of controller for any particular
environment is a very systematic procedure. There are many ways and means that
how a particular type of system may be controlled through which type of controller.
Usually type of controller is selected using only quantitative considerations
stemming from the analysis of the system and ending at the properties of that
particular controller and the control objective. Proportional, Integral and Derivative
control modes also affect the response of the system. Following is the summarized
criterion to select that appropriate controller for any process depending upon the
detailed study of the controller and control variable along with process severity.
91
If Possible Use a Simple Proportional Controller
Simple P-controller can be used if we can achieve acceptable off-set with
not too high values of gain. So for gas pressure and liquid level control, usually a
simple proportional controller may be used.
If a Simple P-Controller is not Suitable, Use PI-Controller
A steady-state error always remains for proportional controller so in
systems where this off-set is to be minimized, a PI-controller is incorporated. So in
flow control applications, usually PI-controller is found.
Use PID-Controller to Increase the Speed of the Closed Loop Response
The anticipatory characteristics of the derivative control enables to use
somewhat higher values of proportional gains so that off-set is minimized with
lesser derivations and good response of the system. Also it adds the stability to the
system. So this type of control is used for sluggish multi-capacity processes like to
control temperature and composition.
92
7.8. Control Loops
For instrumentation and control of different sections and equipment’s of
plants, following control loops are most often used.
I. Feed backward control loop
II. Feed forward control loop
III. Ratio control loop
IV. Auctioneering control loop
V. Split range control loop
VI. Cascade control loop
I. Feed Backward Control Loop
Feedback control is a control mechanism that uses information from
measurements to manipulate a variable to achieve the desired result.
II. Feed Forward Control Loop
Feed-forward control, also called anticipative control, is a control
mechanism that predicts the effects of measured disturbances and takes corrective
action to achieve the desired result.
III. Ratio Control Loop
Ratio control loop is used to maintain the flow rate of one (dependent
controlled feed) stream in a process at a defined or specified proportion relative to
that of another (independent wild feed stream) in order to control the composition
of a resultant mixture.
93
IV. Auctioneering Control Loop
Auctioneering is the process of choosing one output signal from a set of
multiple input signals. In order to use auctioneering in your control process, you
will first need to have multiple signals all measuring the same variable. The signals
will then all be sent to a set of selectors aligned in series. For each selector, there
will be two inputs. For the first selector, the two inputs will be the first two signals
from the device being controlled. For each subsequent selector, one signal will be
the output signal from the previous selector, while the other input signal will be the
next signal from the device.
V. Split Range Control Loop
In this loop controller is per set with different values corresponding to
different action to be taken at different conditions. The advantage of this loop is to
maintain the proper conditions and avoid abnormalities at very different levels.
VI. Cascade Control Loop
This is a control in which two or more control loops are arranged so that the
output of, one controlling element adjusts the point of another controlling element.
This control loop is used where proper and quick control is difficult by simple feed
forward or feed backward control. Normally first loop is a feed control loop. We
have selected a cascade control loop for our heat exchanger in order to get quick
on proper control.
94
7.9. Control Schemes of Gasifier
The gasifiers of the past were crude, inconvenient devices. Today's gasifiers
are evolving toward safer, automated processes that make use of a wide range of
present-day instruments and controls.
Thermocouples (such as chromel-alumel type K) should be used to measure
various gasifier temperatures, especially below the grate, as a check for normal or
abnormal operation. Temperatures at the grate should not exceed 950 0K higher
temperatures indicate abnormal function. Consequently, the signal from the
thermocouple can be used by a control system or an alarm system.
For gasification reaction, gasification agent steam is required which must
be enter into gasifier at specific temperature and pressure. Therefore temperature
and pressure of steam must be controlled by the used of thermocouple and pressure
gauge respectively.
Figure 7.1: Control Schemes of Gasifier
95
7.10. Control Schemes of Scrubber
Pressure of entering streams will be sensed by the differential pressure sensor
and temperature of entering streams and exits stream must be noticed by the used
of thermocouple as well as for maximum efficiency of scrubber flow rate of the
water entering the scrubber must be control by the used of flow rate meter and
temperature is controlled by thermocouple.
Figure 7.2: Control Schemes of Scrubber
CHAPTER 8
COST ESTIMATION
8.1. Introduction
Before the plant to be operated, specified money must be supplied to
purchase and install the equipment. The capital needed to supply the necessary
plant facilities is called fixed capital investment while that for the operation of the
plant is called the working capital and sum of two capitals is called total capital
investment.
It is essential that chemical engineer be aware of the many different types
of cost involved in manufacturing processes. Capital must be allocated for direct
plant expenses; such as those for raw materials, labor, and equipment. Besides
direct expenses, many other indirect expenses are incurred, and these must be
included if a complete analysis of the total cost is to be obtained.
97
8.2. Fixed Capital Investment
Manufacturing fixed capital investment represents the capital necessary for
the installed process equipment. With all auxiliaries that are needed for the
complete process operation. Expenses for piping, instruments, installation, and
foundation and site preparation are typical examples of cost included in the
manufacturing fixed capital investment. Fixed capital required for the construction
overhead and for all plant components that are not directly related to the process
operation is designed as the non-manufacturing fixed capital investment.
8.3. Working Capital
The working capital for an industrial plant consists of the total amount of
the money invested in:
Raw material and supplies carried in stock.
Finished products in stock and semi-finished products in the process of
being manufactured.
Accounts receivable.
Cash kept on hand for monthly payments of operating expenses, such as
salaries, wages and raw material purchases.
Accounts payable
Taxes payable
8.4. Cost Index
Is an index value for a given point in time showing the cost at that time
relative to certain base time. If the cost at any time in past is known, the equivalent
cost at the present time can be determined by multiplying the original cost by the
ratio of present of index value to the index value applicable when the original cost
was obtained.
Present cost = Original cost × cost index value at present time/cost index value
at past.
98
8.5. Purchased Equipment Cost
8.5.1. Estimate of Dryer Cost [38]
Surface area of dryer = 2πr (r + h) = 35.2 m2 = 378.86 ft2
From figure 16.29 [41]
Cost of Rotary dryer = $ 55000
CE Plant Cost Index= 356, January 1990
CE Plant Cost Index = 576.2, June 2014
Total Cost of Dryer in 1990 = $ 55000
Total Cost of Dryer in 2014 = (55000)(576.2/356)
Total Cost of Dryer in 2014 = $ 89019.66
8.5.2. Estimate of Ball Mill Cost [39]
Ball mill: C = 50.0 W 0.69, 1 < W< 30 tons/hr
W = 0.6836 ton/hr
For 2 ton / hr
C = 80664.17 $
By Six-tenths-factor rule (Slope 0.6) [38]
Cost of Ball mill (for the required capacity) = 42358.82 $
99
CE Plant Cost Index= 325, Middle 1985
CE Plant Cost Index = 576.2, June 2014
Total Cost of Ball mill in 2014 = (42358.82)(576.2/325)
Total Cost of Ball Mill in 2014 = $ 75098.94
8.5.3. Estimate of Gasifier Cost [38]
Thickness = 𝑃∗𝑅𝑖
𝑆𝐸𝑗−0.6𝑃+ 𝐶𝑐
P = Max allowable internal pressure = 1.2 atm = (20% Design pressure)
Ri = Inner radius= 0.52 m = 1.71 ft
Ej = Efficiency of joint = 0.85
S = Max allowable working stress = 12000 psi = 816.55 atm
Cc = Corrosion allowance = 0.125 in
By putting the values
Thickness = 0.5 in
Weight of Shell = π*ID*H*Thickness*Density
Density = 490.06 lb/ft3
By putting the values
Weight of Shell = 3804.04 lb
Weight of two heads = 2π[(D/2)^2]*Thickness * Density
= 373.22 lb
Total weight including 20% increase = 5012.712 lb
100
Cost = 80W-0.34
= 22137.98 $
Cost factor of stainless steel = 2.5
Total cost = 55344.95 $ (January 1990)
CE Plant Cost Index= 356, Jan 1990
CE Plant Cost Index = 576.2, June 2014
Present cost = $ 89546.89
Cost of distributor plate (2004) = 544 $ [36]
Present cost of distributor plate = 777.79 $
Total Cost of Gasifier in 2014 = $ 90324.69
8.5.4. Estimate of Cyclone Separator Cost [39]
Standard duty Cyclone: C = O.65Q0.91, 2 < Q < 40 KSCFM
Q = 2.351 K SCFM
Cost of Cyclone (for the required capacity) = 1414.9 $
CE Plant Cost Index= 325, Middle 1985
CE Plant Cost Index = 576.2, June 2014
Total Cost of Cyclone in 2014 = (1414.9)(576.2/325)
101
Total Cost of Cyclone in 2014 = $ 2508.67
8.5.5. Estimate of Scrubber Cost [39]
For Packed Towers;
C = f1Cb + VpCp + Cp1
Pressure = 1.013 bar
The design pressure of the vessel should be 10% above the operating pressure
Column Diameter (Di) = 0.9 m
Vessel Length (L) = 6.25 m
Height of packing = 4.25 m
Shell Thickness = 10mm
Column Diameter (Do) = 0.92 m
Density of Stainless Steel (ρss) = 7850 kg/m3
Mass of Top and Bottom Heads = 147 kg each
𝑀𝑎𝑠𝑠 𝑜𝑓 𝐶𝑜𝑙𝑢𝑚𝑛 = (𝜋(𝐷𝑜2 − 𝐷𝑖2)/4)(𝐿)(ρ𝑠𝑠) + 𝑀(𝑡𝑜𝑝 ℎ𝑒𝑎𝑑) +
𝑀(𝑏𝑜𝑡𝑡𝑜𝑚 ℎ𝑒𝑎𝑑)
Mass of Column = 1696.62 kg = 3732.57 lb
Cb = exp[6.629+ 0.1826(1n W) + 0.02297(ln W)2]
Cb = $ 16465.13
Cp1 = 246.4D0.7396L0.7068
Cp1 = $ 4639.66
F1 = 1.7 (For stainless steel)
Vp (packing volume) = Area x Height
Vp (packing volume) = 95.48 ft3
Cp = 16.6 $/ft3 (for 1.5 in Ceramic Intalox Saddles)
C = f1Cb + VpCp + Cp1
By putting the values
102
C = $ 34215.35
CE Plant Cost Index= 325, Middle 1985
CE Plant Cost Index = 576.2, June 2014
Total Cost Scrubber in 1985 = $ 34215.35
Total Cost Scrubber in 2014 = (34215.35)(576.2/325)
Total Cost Scrubber in 2014 = $ 60661.18
Table 8.1: Total Purchased Cost of Major Equipments
Equipment Purchased cost $
Dryer 89019.66
Ball Mill 75098.94
Gasifier 90325.69
Cyclone Separator 10034.68
Scrubber 60661.18
Total $ 325140.15
103
Estimation of Fixed Capital [36]
Table 8.2: Typical factors for the estimation of project fixed capital cost
Physical Plant cost PPC= PCE* (1+f1+f2+….f9)
= 325140.15 (1+0.45+0.45+0.15+0.10+0.10+0.45+0.20+0.05+0.20)
PPC = $1,024,191.473
Table 8.3: Fixed Capital Cost
Fixed capital cost =PPC*(1+F10+F11+F12)
Fixed capital cost = $1,024,191.473 (1+0.25+0.05+0.10)
Fixed capital cost = $1,433,868.06
Items Factors
Equipment erection F1 0 .45
Piping F2 0.45
Instrumentation F3 0.15
Electrical F4 0.10
Buildings, process F5 0.10
Utilities F6 0.45
Storage F7 0.20
Site development F8 0.05
Ancillary buildings F9 0.20
Items Factors
Design and Engineering F10 0 .25
Contractor’s fee F11 0.05
Contingency F12 0.10
104
8.6. Estimation of Working Capital
Working capital = 10% of fixed capital to cover the cost of the fixed capital used
= $1,433,868.06*0.1
= $ 143,386.81
8.7. Total Investments
Total investments required = Fixed capital + Working capital
= $1,433,868.06 + $ 143,386.81
= $ 1577254.87
8.8. Production Costs [36]
I. Variable costs
1. Raw materials
= $ 716335.4 /year
2. Miscellaneous materials
= 10 % of maintenance cost
= $ 14338.68 /year
3. Utilities
= 5% of maintenance cost
= $ 7169.34 /year
Total Variable Cost = $ 737843.42
105
II. Fixed costs
Maintenance Cost
= 10 % of fixed capital
= $ 143,386.81
Operating labor
= 15 % of the total operating cost.
= $ 236,588.23
Laboratory costs
= 20 % of Operating Labor
= $ 47,317.65
Supervision
= 20 % of Operating Labor
= $ 47,317.65
Plant overheads
= 50 % of Operating Labor
= $ 118,294.11
Capital charges
= 10 % of the fixed capital
= $ 143,386
Insurance 1 per cent of the fixed capital
= 1 % of the fixed capital
= $ 14,338.6
12. Local taxes 2 per cent of the fixed capital
= 2 % of the fixed capital
= $ 28,677.2
106
13. Royalties 1 per cent of the fixed capital
= 1 % of the fixed capital
= $ 14,338.6
Total Fixed Cost = $ 793,644.85
Direct production costs = $ 1,531,488.27
Sales expense + General overheads
= 20% of the direct production cost
= $ 306297.65
Annual production cost = $ 1,837,785.92
Production cost $/kg = 𝐴𝑛𝑛𝑢𝑎𝑙 𝑃𝑟𝑜𝑑𝑢𝑐𝑡𝑖𝑜𝑛 𝐶𝑜𝑠𝑡
𝐴𝑛𝑛𝑢𝑎𝑙 𝑃𝑟𝑜𝑑𝑢𝑐𝑡𝑖𝑜𝑛 𝑅𝑎𝑡𝑒
= 0.35 $/Kg
CHAPTER 9
HAZOP STUDY
9.2. Introduction
The Hazard and Operability Study (or HAZOP Study) is a standard hazard
analysis technique used in the preliminary safety assessment of new systems or
modifications to existing ones. The HAZOP study is a detailed examination, by a
group of specialists, of components within a system to determine what would
happen if that component were to operate outside its normal design mode. The
effects of such behavior is then assessed and noted down on study forms. The
categories of information entered on these forms can vary from industry to industry
and from company to company.
9.3. Objectives of HAZOP Study
To identify the potential risks
To identify and study features of the design that influence the probability of
a hazardous incident occurring.
To familiarize the study team with the design information available.
108
To ensure that a systematic study is made of the areas of significant hazard
potential.
9.4. Keywords used in HAZOP STUDY
Keywords are used to focus the attention of the team upon deviations and
their possible causes. These keywords are divided into two sub-sets:
Primary Keywords which focus attention upon a particular aspect of the
design intent or an associated process condition or parameter.
Secondary Keywords which, when combined with a primary keyword,
suggest possible deviations.
9.5. Primary Keywords
These reflect both the process design intent and operational aspects of the
plant being studied. Typical process oriented words might be as follows. The list
below is purely illustrative, as the words employed in a review will depend upon
the plant being studied.
Table 2: Primary Keywords
Flow Temperature
Pressure Level
Separate (settle, filter, centrifuge) Composition
React Mix
Reduce (grind, crush, etc.) Absorb
Corrode Erode
109
9.6. Secondary Keywords
Secondary keywords when applied in conjunction with a Primary Keyword,
these suggest potential deviations or problems. They tend to be a standard set as
listed below:
Table 3: Secondary Keywords
Guide Words Meaning
No
Less
More
Part of
As well as
Reverse
Other than
Negation of design intent
Quantitative decrease
Quantitative increase
Qualitative decrease
Qualitative Increase
Logical opposite of the intent
Complete substitution
9.7. How to Conduct a Hazop Study
HAZOP study is conducted in following steps:
1. Specify the purpose, objective, and scope of the study. The purpose may be
the analysis of a new plant or a review of the risk of unexisting unit. Given
the purpose and the circumstances of the study, the objectives listed above
can more specific.
Select the HAZOP study team. The team who will conduct the Hazop study
should consist of personnel with a good understanding of the process and
plant to be reviewed. The group should ideally contain about six members,
with perhaps an absolute upper limit being set at nine. In a study in which
both contractor and client are participating, it is desirable to maintain a
balance between the two in terms of team membership so that neither side
feels outnumbered.
2. Make a preparatory work. It is most important that, before a study
commences, work that can be conveniently done beforehand is carried
110
out. This is not only essential in some respects for the proper structuring of
the study and the team, but will also greatly increase the efficiency of the
Hazop and thus retain the interest and enthusiasm of the participants.
3. This preparatory work will be the responsibility of the Chairman, and the
requirements can be summarized as follows:
Assemble the data
Understand the subject
Subdivide the plant and plan the sequence
Mark-up the drawings
Devise a list of appropriate Keywords
Prepare Node Headings and an Agenda
Prepare a timetable.
Select the team.
111
9.8. HAZOP Method Flow Diagram
Figure 9.1: HAZOP Method Flow Diagram
112
9.9. Hazop Analysis on Fluidized Bed Reactor
Node: Fluidized Bed Reactor
Parameter: Temperature of bed
Hazards and Operability Study Worksheet
HAZARD AND OPERABILITY STUDY
NODE: FLUIDIZED BED REACTOR
PROCESS PARAMETER: FLOW OF STEAM
No. Guide
Word
Cause Consequences Actions
1. No Boiler is not
working.
Gasifier will not
work.
Boiler should
work
properly.
2. Less Less water flow
rate inside the
boiler / Boiler
Inefficiency
Gasifier will not
work properly.
(Incomplete
reaction)
Flow rate of
steam must
be
maintained.
3. More Steam flow rate
is not controlled.
Product not
obtained in desired
form.
Steam inlet
flow rate
must be
controlled
4. Other than Minerals present
in water.
Formation of
undesired products
Water must
be treated
properly
113
HAZARD AND OPERABILITY STUDY
NODE: FLUIDIZED BED REACTOR
PROCESS PARAMETER: TEMPERATURE OF BED
No. Guide
Word
Cause Consequences Safeguards
1. Less Inefficiency of
boiler.
Gasifier will not
work properly.
Boiler should
work properly.
2. More Poor
temperature
control system.
Desired yield is
not obtained.
Temperature
of steam must
be maintained.
9.10. Hazop Analysis on Dryer:
Node: Dryer
Parameter: Temperature of Inlet air
HAZARD AND OPERABILITY STUDY
NODE: DRYER
PROCESS PARAMETER: TEMPERATURE OF INLET AIR
No. Guide
Word
Cause Consequences Safeguards
1. Less Air heater is not
working
properly
Proper drying will
not take place.
Steam heater
must be
repaired.
2. More Temperature of
air is not
controlled
Combustion of
Coal and
Sugarcane
bagasse.
Air inlet
temperature
must be
controlled
114
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115
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117
APPENDIX A
A1: Char Gasification Reaction Equilibrium with Temperature at Pressure
of 1 atm.
A2: Equilibrium Methane Conversion at S/C =2
118
A3: CO Conversion in Water Gas Shift Reaction
A4: Temperature Pattern in Dryer
119
A5: Cyclone Pressure Drop Factor
A6: Standard Cyclone Dimension, High Gas Rate Cyclone
120
A7: Performance Curves, High Gas Rate Cyclone
A8: Scaled Performance Curve Cyclone
121
A9: Generalized Pressure Drop Correlation
A10: Number of Transfer units NOG as a function of y1/y2 with mGm/Lm as
Parameter
122
A11: Installed Cost of Dryer
A12: Columns Plates. Time Base mid-2004
Installed Cost = (Cost From Figure) * Material Factor