DESIGN OF A PLANT FOR THE PRODUCTION OF 750,000 STANDARD M³/DAY OF HYDROGEN BY PARTIAL OXIDATION OF HEAVY OIL FEEDSTOCK

ABSTRACT

The design of a plant producing hydrogen by partial oxidation of oil feedstock has been carried out. Purity of hydrogen obtained from the process was set at 98%, with minimal levels of impurities. The main process involved include the partial oxidation of the heavy fuel oil feedstock in a flame lined reactor called a gasifier to produce hydrogen. The successive units in the plants are purification units to ensure that the product specification of 98% purity of hydrogen is met. A water gas shift converter was also present to improve hydrogen yield. Waste heat energy generated from the gasification process is used in the waste heat boiler to generate steam for the plant. These purification units include: Quencher, Initial H₂S scrubber, final H₂S removal, CO₂absorption.

The chemical and mechanical design of the main equipments in the plants was carried out and the results presented. Costing and HAZOP analysis of each item of equipment is also provided. Detailed economic and profitability analysis was done and general safety and environmental issues were addressed.

TABLE OF CONTENTS

CONTENTS                 PAGE

CERTIFICATION

DECLARATION

DEDICATION

LETTER OF TRANSMITAL

ACKNOWLEDGEMENT v

ABSTRACT vi

LIST OF TABLES

LIST OF FIGURES xix

LIST OF SYMBOLS xxi

CHAPTER ONE

1 INTRODUCTION

1.1 HYDROGEN GAS: BRIEF HISTORY AND MARKET VALUE

1.2 THE HYDROGEN ECONOMY

1.3 LIMITATIONS OF HYDROGEN PRODUCTION

CHAPTER TWO

2 LITERATURE REVIEW

2.1 THE GASIFIER

2.1.1 FLUIDISED BED GASIFIERS

2.1.2 MOVING BED GASIFIERS

2.1.3 ENTRAINED FLOW GASIFIERS

2.2 THE WASTE HEAT BOILER

2.2.1 TYPES OF WASTE HEAT BOILERS

2.2.2 WASTE HEAT BOILER DESIGN CONSIDERATIONS

2.3 QUENCHING

2.3.1 HORIZONTAL DESIGN: CFQ QUENCHER

2.3.2 VERTICAL DESIGN: CCQ-QUENCHER

2.3.3 DESCRIPTION OF DESIGN

2.4 SCRUBBING SYSTEMS (INITIAL H2S REMOVAL)

2.4.1 TYPES OF SCRUBBERS

2.4.2 SCRUBBING MECHANISM

2.4.3 SOLVENTS FOR SCRUBBING

2.4.4 ABSORBER DESIGN

2.5 CO CONVERSION STAGE

2.5.1 DESIGN AND OPERATION OF WATER GAS SHIFT REACTORS

2.6 SATURATOR/DESATURATOR. 30

2.6.1 THE SATURATOR 30

2.6.2 THE DESATURATOR 31

2.7 FINAL H2S REMOVAL (CHEMOSORPTION)

2.7.1 PRESSURE SWING ABSORBER DESIGN

2.8 CARBON DIOXIDE REMOVAL

2.8.1 THE UNIT OPERATION

2.8.2 EQUIPMENT

2.8.3 HOT POTASSIUM CARBONATE (BENFIELD)  FOR CO2 REMOVAL

2.9 ANCILLIARY EQUIPMENT 39

2.9.1 RAM PUMPS 39

2.9.2 HEAT EXCHANGERS  42

2.9.3 AIR-COOLED EXCHANGERS                       43

2.9.4 FINNED TUBES 44

2.9.5 PLATE HEAT EXCHANGERS 45

2.9.6 SHELL AND TUBE HEAT EXCHANGERS 47

2.9.7 MEAN TEMPERATURE DIFFERENCE ( DRIVING FORCE) 48

CHAPTER THREE

3 MATERIAL AND ENERGY BALANCE

3.1 MATERIAL BALANCE

3.2 ENERGY BALANCE AROUND CO CONVERSION STAGE

3.3 ENERGY BALANCE AROUND THE GASIFIER 58

CHAPTER FOUR                                                                                                                   61

4 CHEMICAL ENGINEERING DESIGN 61

4.1 TANK ONE (FEED STORAGE TANK FARMS) 61

4.2 TANK 2 (FRESH FEED SURGE DRUM) 62

4.3 TANK 3 (STORAGE TANK FARM) 62

4.4 THE HEAVY FUEL OIL GASIFIER 64

4.5 THE WASTE HEAT BOILER 67

4.6 THE QUENCHER 72

4.6.1 DESIGN PROCEDURES 73

4.6.2 TYPES OF PACKING 73

4.6.3 HEAT CAPACITY 74

4.6.4 COLUMN DIAMETER 76

4.6.5 HEIGHT OF THE COLUMN: Z 76

4.7 INITIAL  H2S REMOVAL 77

4.7.1 ASSUMPTIONS 77

4.8 CARBON MONOXIDE CONVERSION UNIT 79

4.8.1 REACTOR TYPE: FIXED BED REACTOR 80

4.8.2 CATALYST: CHROMIUM – PROMOTED IRON OXIDE 81

4.8.3 CATALYTIC CONVERTER SIZING 84

4.9 FINAL H2S REMOVAL STAGE 85

4.9.1 NORMAL OPERATION 86

4.9.2 REGENERATION OPERATION 86

4.9.3 THE EQUIPMENT 87

4.9.4 ASSUMPTION 89

4.9.5 REGENERATION 91

4.10 CARBON DIOXIDE REMOVAL 92

4.10.1PACKING MATERIAL: PALL RINGS (PLASTIC) 94

4.11 HEAT EXCHANGER DESIGN 96

4.11.1  THE PROPOSED DESIGN 96

4.11.2  OPTIMISATION 97

4.11.3  VISCOSITY CORRECTION FACTOR 97

4.11.4 OXYGEN STREAM PRE-HEATER. 98

4.11.5 AIR-COOLED EXCHANGERS 102

4.12    PUMP SIZING 108

4.12.1APPROXIMATE PUMP-SIZING CALCULATION PROCEDURE 109

4.12.2THE METERING RAM PUMP 111

CHAPTER FIVE 115

5 MECHANICAL DESIGN 115

5.1 THE GASIFIER 115

5.1.1 DESIGN PRESSURE 115

5.1.2 DESIGN TEMPERATURE 115

5.1.3 MATERIAL OF CONSTRUCTION 115

5.1.4 THICKNESS OF THE REFRACTORY BRICK 116

5.1.5 MINIMUM WALL THICKNESS 116

5.1.6 MINIMUM HEAD THICKNESS 117

5.1.7 VESSEL SUPPORT 117

5.2 THE WASTE HEAT BOILER 117

5.3 THE QUENCHER 118

5.3.1 MATERIALS FOR CONSTRUCTION 118

5.3.2 PACKING 118

5.3.3 DESIGN PRESSURE 118

5.3.4 DESIGN TEMPERATURE 119

5.3.5 TENSILE STRENGTH 119

5.3.6 DESIGN STRESS 120

5.3.7 CHOICE OF CLOSURE 120

5.4 THE INITIAL H2S ABSORBER 121

5.4.1 SHELL THICKNESS 121

5.4.2 SUPPORT FOR ABSORBER 124

5.5 THE CO CONVERTER 125

5.5.1 OPERATING AND DESIGN TEMPERATURES AND PRESSURES 125

5.5.2 VESSEL DIMENSIONS AND ORIENTATION 125

5.5.3 CONSTRUCTION MATERIALS 126

5.5.4 TYPE OF VESSEL HEAD 126

5.5.5 OPENINGS REQUIRED 127

5.5.6 HEATING AND COOLING JACKETS OR COILS 127

5.5.7 INTERNAL FITTINGS SPECIFICATION 127

5.6 THE FINAL H2S ABSORBER 128

5.6.1 DESIGN PRESSURE 128

5.6.2 WALL THICKNESS 129

5.6.3 SUPPORT 129

5.7 THE CO2 ABSORBER 129

5.7.1 SHELL THICKNESS 129

5.7.2 AXIAL STRESS DUE TO PRESSURE 130

5.7.3 STRESS DUE TO DEAD LOAD 130

5.7.4 EFFECT OF ATTACHMENTS 132

5.7.5 SUPPORT FOR ABSORBER 133

5.8 THE HEAT EXCHANGERS 133

5.8.1 FIRST EXCHANGER 134

CHAPTER SIX 138

6 PLANT HAZARD AND OPERABILITY STUDY 138

6.1 GASIFIER 138

6.2 WASTE HEAT BOILER 139

6.3 HEAT EXCANGERS 140

6.4 INITIAL H2S REMOVAL 141

6.5 SATURATOR/DESATURATOR 145

6.6 CO2 REMOVAL 147

6.7 STORAGE TANK 149

6.8 SAFETY 150

6.8.1 STORAGE TANK 150

6.8.2 THE HEAT EXCHANGERS 150

6.8.3 THE CO CONVERTER 151

6.8.4 SATURATOR – DESATURATOR 152

6.8.5 THE FINAL H2S ABSORBER 153

CHAPTER SEVEN 158

7 COSTING 158

7.1 ECONOMIC AND PROFITABILITY ANALYSIS: GASIFIER 158

7.2 ECONOMIC AND PROFITABILITY ANALYSIS: QUENCHER 161

7.3 COSTING (CO CONVERSION) 163

7.4 COST AND FEASIBILITY ANALYSIS (FINAL H2S REMOVAL) 164

7.5 COSTING AND PROFITABILITY (CO2 REMOVAL) 166

7.6 ESTIMATION OF COST (1ST HEAT EXCHANGER) 168

7.7 ESTIMATION OF COST (2ND HEAT EXCHANGER) 169

CHAPTER EIGHT                                                                                                                 170

8     CONCLUSION AND RECOMMENDATION                                                              170

8.1  CONCLUSION                                                                                                               170

8.2  RECOMMENDATION                                                                                                  171

REFERENCES          172

A.1 SIMULATION          175

A.2 PLANT AND SITE LAYOUT        176

A.3 PROCESS INFORMATION        177

A.4 CALCULATIONS                                                     178

LIST OF TABLES

TABLE	TITLE	PAGE
2.1	Classification of Gasifiers	10
2.2	Comparison of Calculated and Experimental Solubility Data for Scrubbing Solvents	23
3.1	Composition before H2S Removal	51
3.2	Composition after H2S Removal	51
3.3	Composition Out of WGS Reactor	53
3.4	Final H2S Remover Outlet	54
3.5	Percentage Composition of Constituents from Final H2S Removal Outlet	54
3.6	Composition of the Inlet into the CO Conversion Stage	57
3.7	Composition of the Outlet from the CO Conversion Stage	57
3.8	Energy Balance of Outlet Stream from Gasifier	59
3.9	Energy Balance of Inlet Stream to Gasifier	60
4.1	Tank 1 Dimensions	61
4.2	Tank 2 Dimensions	62
4.3	Tankage Dimensions	63
4.4	Rate Constant Values At Specific Temperatures	66
4.5	Equilibrium Constants at Specific Temperatures	66
4.6	Composition of Crude Gas Exiting The Gasifier (Wet Basis)	69
4.7	Composition of Inlet Stream Into Quencher	74
4.8	Composition of Outlet Stream from Quencher	75
4.9	Equilibrium Data for CO at Different Temperatures	80
4.10	Catalyst Properties of Chromium Promoted Iron Catalyst	82
4.11	Reactor Dimensions of Catalytic Converter	84
4.12	Composition of Inlet and Outlet Streams for CO2 Removal	93
4.13	Equilibrium Data for CO2 Removal	94
4.14	Summary of CO2 Absorber Dimensions	95
4.15	Summary of Proposed Heat Exchanger Design	96
4.16	Heat Transfer Coefficients with Corresponding Area And Face Velocity Values	105
4.17	Corresponding Face-Area and Face-Velocity Values	106
4.18	Approximate Frictional Pressure Drop Across Process	111
5.1	Tensile Strength Values At Different Temperatures	119
5.2	Design Stress Values at Different Temperatures	120
6.1	Hazop Analysis for Gasifier	138
6.2	Hazop Analysis or Waste Heat Boiler	139
6.3	Hazop Analysis for Heat Exchangers	140
6.4	Hazop Analysis for Initial H2s Removal	141
6.5A	Hazop Analysis for Saturator	145
6.5B	Hazop Analysis for Desaturator	146
6.5C	Hazop Analysis for Lines 1 And 2	146
6.6A	Hazop Analysis for Co2 Removal Inlet	147
6.6B	Hazop Analysis for Co2 Removal Outlet	148
6.7	Hazop Analysis for Storage Tanks	149
6.8	Flammability Limits of Different Compounds	155
7.1	Inflation Rates By Month And Year (1999-2011), US Inflation Calculator	159
7.2	Refractory Brick Density Calculation	160
7.3	Inflation Rates By Month And Year (2005-2011) (For Bare Vessel)	161
7.4	Inflation Rates By Month And Year (2005-2011) (For Packing Material)	162
7.5	Inflation Rates By Year (2005-2011) ( For H2s Removal Vessel)	165

LIST OF FIGURES

FIGURE	TITLE	PAGE
2.1	Shell Oil Gasifier	12
2.2	The Quenching Unit As Conditioned By Design Specifications	18
2.3	Schematic Diagram of a Packed Column	19
2.4	Structure of Sterically Hindered Amines	23
2.5	Adiabatic Stages in Water Gas Shift Reaction	26
2.6	Schematic Diagram Of Final H2s Removal Stage	32
2.7	Benfield Process For Co2 Removal	38
2.8	Ram Pump Schematics	39
2.9	Ram Pump Schematics	40
2.10	Ram Pump	41
2.11	Air-Cooled Exchangers	44
2.12	Finned Tubes	45
2.13	A Plate Heat Exchanger	46
2.14	Assemblage Of Baffles And Tubes Inside A Shell And Tube Heat Exchanger	48
2.15	Counter-Current Flow In A Heat Exchanger And The Temperature Profile Within The Exchanger	48
2.16	Co-current Flow Through A Heat Exchanger With The Temperature Profile Through The Equipment	49
4.1	The Heavy Fuel Oil Gasifier	64
4.2	The Quencher	72
4.3	The Initial H2S Scrubber	77
4.4	The CO Catalytic Converter	79
4.5	The Final H2S Remover	85
4.6	CO2 Removal	92
4.7	Air-Cooled Exchangers	103
4.8	Finned Tubes	104
5.1	Skirt Support	128
5.2	Double Plate With Gusset	128
5.3	Physical Layout Of A 2-Tube Pass Shell And Tube Exchanger	137
7.1	Cost Estimation Data For Shell And Tube Heat Exchangers	168

LIST OF SYMBOLS

2S left in output stream

B = amount of H₂S removed from aP₃

We are told that the H₂S is scrubbed to 15ppm

aP₃  = 0.0201K_mol/hr

so that

0.5 = 0.02011 + b

B = 0.4799 Kmol/hr

TABLE 3.1: COMPOSITION BEFORE H₂S REMOVAL

components	Kmol	Kmol %
H2	47.6	671.636
Co	42.1	594.031
Co2	8.3	117.113
N2	1.40	19.754
CH4	0.1	1.411
H2S	0.5	7.055
		1411

 

TABLE 3.2: COMPOSITION AFTER H₂S REMOVAL

components	Kmol	Kmol %
H2	671.636	47.839
Co	594.031	42.311
Co2	117.113	8.342
N2	19.754	1.401
CH4	1.411	00.1005011
H2S	0.02011	0.001432
	1403.9651	100.0000

From here, the scrubbed gas moves on to the water gas shift reactor, where CO is used up and H₂, CO₂ are produced.

Now, onto the CO conversion stage. The reaction stoichiometry ofWater gas shift is:

CO + H₂O = CO₂ + H₂

Assuming 99.8% conversion was obtained from the simulation we can obtain the composition of the output from the 2^nd stage CO conversion:

Overall conversion = 

0.998 = 

Where aP₅ is the molar flow of  CO out of the WGS reactor.

aP₅ = 1.881Kmol

Amount of CO reacted = 594.031 – 1.881

= 592.8429 Kmol

Also, using the reaction stoichiometry,

Amount of CO produced = 592.8429 Kmol

Amount of H₂ produced = 592.8429 Kmol

Amount of H₂O reacted = 592.8429 Kmol**

** is the total amount of water reacted. From the problem statement, we now that this is made up of water from the gasification process from the saturation and from the water fed into the WGS Reactor.

An oxygen atom balance around the Reactorshould tell exactly how much of this water came from the gasified and this will be done later.

TABLE 3.3: COMPOSITION OUT OF THE WGS REACTOR

Components	Kmol	Kmol %
H2	1264.4789	63.3335
CO	1.1881	0.05951
CO2	709.9559	35.5593
N2	19.754	0.9894
CH4	1.411	0.0707
H2S	0.02011	0.001007
	1996.541	100.001

From here, the gas mixture moves to the final H₂S removal stage, where H₂S is scrubbed to less this 1ppm. Assuming scrubbing of H₂S to 0.8ppm,

 

 

 

TABLE 3.4: FINAL H₂S REMOVAL OUTLET

Components	Kmol
H2	1264.4789
CO	1.1889
CO2	709.9559
N2	19.754
CH4	1.411
H2S	1.06583 x 10-3Kmol

We are required to produce 750,000 stdm³/day of feed i.e. 31250 stdm³/hr. Using the ideal gas relation, this is equivalent to 1277.92 Kmol /hr

Using this, we can calculate the % composition of all constituent except C_o2, since there are molar flows are already fixed

Table 3.5: PERCENTAGE COMPOSITION OF CONSTITUENTS FROM FINAL H2S REMOVAL OUTLET

H2	98.00%
Co	1.1889%
N2	1.55%
CH4	0.1104%
H2S	0.0000834%

CO₂ = 100 – (98.00 + 0.093 + 1.55 + 0.1104 + 0.0000834)

= 0.247%

Molar flow of CO₂ =  x 1277.92

= 3.1503 Kmol /hr

To obtain feed flow rate:

Moles of C in quenched gas

= 591.636 () + 117.113 () + 1.411 ()

=287.842 Kmol of carbon

= 3454.104 kg

The mass of C calculated above amounts to 98.5% of the total carbon in the feed. Using this information, we can calculate the total weight of the Carbon in feed

0.985 C = 3454.104

C = 3506.705Kg

Now, we can calculate the total weight of feed:

85 wt. % of feed is carbon

0.85 (F) = 3506.705kg

F = 4125.535kg/hr feed

We can also calculate the flow rate of steam and air:

Mass of steam = 0.75 (1125.535)

= 3094.1511kg/hr

Now,

4125.535kg =  +  + 

=  751.1912 kmol/hr

1.16kg of O₂ = 0.03625 mole of O₂

1.16kg of O₂ = 0.03625 mol of O₂/0.1821 mol feed

So that:

Total mole of (impure) feed = 

= 149.551 mol of impure O₂

Actual mole of O₂ fed = 0.95 (149.551)

= 142.073 moles of O₂

= 4546.34kg of O₂

3.2 ENERGY BALANCE AROUND CO CONVERSION STAGE

TABLE 3.6 COMPOSITION OF THE INLET INTO THE CO CONVERSION STAGE

COMPONENT	COMPOSITION (Kmol)	Tref (K)	Href (J/gmol)	H380	∆H	MiHi
H	671.636	298	718	10977.2	10259.2	6890448.051
CO	594.031	298	728	11243.5	10515.5	6246532.981
C02	117.113	298	912	16200.5	15288.5	1790482.101
N2	19.254	298	728	11099.5	10371.5	199692.861
CH4	1.411	298	879	16800.5	15921.5	22465.237
H2S	0.02011	298	845	13522.3	12677.3	254.941
H2O	553438	298	837	13072.2	12235.2	653824617.6
TOTAL						668.97 ˣ 106

TABLE 3.7 COMPOSITION OF THE OUTLET FROM THE CO CONVERSION STAGE

COMPONENT	COMPOSITION (Kmol)	Tref (K)	Href (J/gmol)	H380	∆H	MiHi
H	1264.4789	298	718	10988.5	10278.5	12986830.54
CO	1.1888	298	728	11206.5	10478.5	12449.505
C02	702.9559	298	912	16231.5	15319.5	10768.93291
N2	19.754	298	728	11139.5	10411.5	205668.771
CH4	1.411	298	879	16819.5	15940.5	22492.0455
H2S	0.02011	298	845	13729	12884	259.097
H2O	52854	298	837	13171	12334	651901.236
TOTAL						675652768.9

 

To obtain the heat of reaction with respect to H₂ produced,

= (1264.4789 – 671.636) × (- 4.2 × 10⁴ KJ/hr Kgmol)

= 5928429Kgmol × 4.2 × 10⁴ KJ/ Kgmol

= – 24.899 × 10⁶ KJ

Q = ₀H₀ – ₁H₁ + ∆H_rn

= (675.652 × 10⁶ – 668.97 × 10⁶) + (- 24.899 × 10⁶)

= – 18.217 × 10⁶ KJ/hr

A total of 18.217 × 10⁶ KJ/hr of heat is given off in the CO conversion unit. This large amount of heat was utilised in the preheating of the unconverted gas in heat exchanger

2.       ENERGY BALANCE AROUND THE GASIFIER

The Geeral energy balance equation reduces to

ASSUMPTIONS

The enthalpy content of the crude gas leaving the gasifier is majorly due to the gasification reaction, which is:

But,

Ignoring all side reactions, the calorific value for the gasification reaction is equal to 4.29 x 10⁷KJ/Kg

The reference temperature of 50% is chosen for convenience, the general energy balance equation becomes;

TABLE 3.8: ENERGY BALANCE OF OUTLET STREAM FROM GASIFIER

Composition	Kgmol	Tref	Href,  J/gmol	H1573  J/gmol		moHo
H2	671.636	323	1485.25	39064.25	37579	25.24 x 106
CO	594.031	323	1514.25	41796.75	40282.5	23.93 x 106
CO2	117.113	323	1965.25	66368.25	64403	7.54 x 106
N2	19.754	323	1513.25	41343.75	39830.5	786.81 x 103
CH4	1.411	323	1897.5	81440.75	79543.25	112.24x 103
H2S	7.055	323	1749.75	54078	52303.25	369.0 x 103
H2O	94.031	323	1749.75	51823.75	50074	29.74 x 106
C		323	–	–	–	87.72 x 106

 

 

TABLE 3.4 ENERGY BALANCE OF INLET STREAM TO GASIFIER

Htrh	Tref	Href,  J/gmol	H1573  J/gmol	J/gmol	miHi	moHo
O2	323	1530.5	68.11	5280.5	1.95 x 106	369.07
N2	323	1513.25	6644	5130.75	99.66 x 103	19.425
H2O	323	1749.75	7752	6002.25	1.03 x 106	171.90
					3.080 x 106

is considered here to be the calorific value of the fuel oil

 = 87.72 

CHAPTER FOUR

4           CHEMICAL ENGINEERING DESIGN

1.       TANK ONE (FEED STORAGE TANK FARMS)

The feed storage tank farm will be designed in such a way that each tank will be able to hold about 5 days of feed in storage. There will be a total of ten feed storage tanks in the tank farm

• The tank is a vertical, cylindrical shaped storage tank with a fixed and floating roof

• Floating roof rises and falls with the liquid level inside the tank, thereby decreasing the vapor space above the liquid level. They are considered a safety requirement as well as a pollution prevention measure.
• A vertical tank is preferred over a horizontal tank since it takes up less space, is easily supported on concrete slabs and allows maximum mixing.
• The storage tanks are supported with a skirt support as they do not impose concentrated loads on the vessel shell
• A drain line is installed so that water at the bottom of the tank (caused by the difference between the internal and external temperature) can be regularly removed.

TABLE 4.1: TANK 1 DIMENSIONS

Top of shell height	100%	23.21m
Design fill level	95%	22.05m
Normal fill level	85%	19.73m
Minimum fill level	35%	8.12m
Bottom of shell height	0%	0

2.       TANK 2 (FRESH FEED SURGE DRUM)

This tank is similar to tank in function, configuration, material of construction and safety devices and structural supports used. The only difference is that it is a smaller tank used to aid the Start-up and shut down procedures. This tank is required to store enough heavy fuel oil for a 3 hour production run.

TABLE 4.2: TANK 2 DIMENSIONS

Top of shell height	100%	6.75m
Design fill level	95%	6.41m
Normal fill level	85%	5.74m
Minimum fill level	35%	2.36m
Bottom of shell height	0%	0m

3.       TANK 3 (STORAGE TANK FARM)

This tank is used for the storage of the final product gas, hydrogen before it is supplied to the immediate consumers. They are designed as spherical tanks since this shape is more ideal for storage of the gaseous product.

TANK DIMENSIONS

For each of the 50 tanks, assume that H = 3D

Similar to tanks 1 and 2, there are safety allowances in the height of the tank.

So that D = 44m

Installed storage capacity is now

TABLE 4.3: TANK STORAGE CAPACITIES

TANK	STORAGE CAPACITY
TANK 1 (FEED STORAGE TANK)	531.25m3
TANK 2 (FRESH FEED SURGE DRUM)	13.281m3
TANK 3 (PRODUCT STORAGE TANK)	75000m3

4.       THE HEAVY FUEL OIL GASIFIER

FIGURE 4.1 THE HEAVY FUEL OIL GASIFIER

The following reactions take place in the gasifier:

    (4.4.1)

  (4.4.2)

        (4.4.3)

        (4.4.4)

       (4.4.5)

The slowest reaction is reaction 4.4.4 (Carbon-steam reaction) and is therefore taken as the rate determining step for the reactor.

C + H₂O → H₂ + CO (rate-determining step)

The rate expression for this reaction is given as:

r = k_v 

Where:

r = rate of reaction (in this case, the rate of disappearance of H₂O)

k_{v =} rate constant

K = equilibrium constant

C_H2O, C_H2 and C_CO = molar concentrations of C_H2O, C_H2 and C_CO respectively (kmol/m³)

Given the table below for rate constant, k_v (from reference source)

TABLE 4.4 RATE CONSTANT VALUES AT SPECIFIC TEMPERATURES

T (K)	952.38	1000
kv (cm3/gmol.s)	1.62	8.6

TABLE 4.5 COMPOSITION OF CRUDE GAS EXITING THE GASIFIER (WET BASIS)

COMPONENTS	AMOUNT (kmol/hr)	COMPOSITION
H2	671.636	0.4232
CO	594.031	0.3743
CO2	117.113	0.0738
CH4	1.411	0.0008
H2O	171.01	0.1077
H2S	7.055	0.0044
N2	19.754	0.0124
C	5.16	0.0033
Total	1587.17	1.0000

Height of the gasifier = 9.64 m

Diameter of the gasifier = 1.99 m

Total molar flow rate of the exit crude gas = 1587.17 kmol/hr

Volumetric flow rate of the exit crude gas =        =3347.88 m³/hr

Residence time, τ = Volume/Volumetric flow rate =  = 0.0089 hr = 32.13s

Space velocity, υ = 1/ τ = 1/32.13s = 0.031 s^-1

 

5.       THE WASTE HEAT BOILER

The waste heat boiler will be designed as a special type of shell and tube heat exchanger called a water tube waste heat boiler, initially mentioned in the previous semester’s work. The water tube design is a very difficult one to obtain but has been selected because of its ability to withstand very large pressure generation like the one specified in this design. Additional energy recovery and fuel conservation can be realized with the use of the water tube boiler, by adding a boiler feed water economizer to the boiler design.

Before we start designing the exchanger, we need to determine the heat duty of the demineralised feed water i.e how much heat it will be required to absorb from the hot flue gas stream.

The calculations leading to the determination of  in the above equation are shown below:

C_p-values are dependent on temperature, so we have to determine the C_p of all the individual components that make up the inlet stream (i.e. at 1300°C)

The heat capacity equations are:

C:        C_p = 11.18 + 1.095 × 10^-2 – 4.891 × 10^5 -2

H₂:       C_p = 

CO:      C_p = 

CO₂:     C_p = 

N₂:       C_p = 

CH₄:     C_p = 

H₂S:     C_p = 

Using these C_p values, we can calculate a bulk C_p for the stream as shown:

TABLE 4.6: BULK C_P VALUES FOR STREAM COMPONENTS

COMPONENT	Cp (J/KMOLOC)	MOLE FRACTION	MOLAR FLOW	IN-STREAM Cp (J/KMOLOC)
C	28.207	0.00362	5.157	0.2089
H2	32.585	0.47191	671.636	13.2328
CO2	58.747	0.08229	117.113	4.1593
CO	35.412	0.41739	594.031	12.7199
N2	35.222	0.01388	19.754	0.4191
CH4	135.761	0.00099	1.411	0.1222
H2S	51.4787	0.00496	6.702	0.0360

From the addition of all the values in the in-stream C_p column, we obtain:

C_p  = 35.92522 J/mol. °C

Substitute the value of C_p  above into Equation (2.1):

From the material balance calculations, we can see that

Inside heat transfer coefficient:

h_i =  1309379.395 W/m²k

Outside heat transfer coefficient

h_o = 2000Btu/ h_o ft^2oF

= 11333.33W/ m^2oC

Overall heat transfer coefficient:

Heat transfer area:

Mean temperature difference:

MTD = 

Number of tubes:

Tube lunght, L:

Tube pitch:  P_i  =
shell diameter: =  1.6646m

Pressure drop (shell side):

Pressure drop (tube side):

Shell wall thickness determination:

P_i= 3000 kPa

D_i = 1.078 m

At the shell side mean temperature of 189.21^oC,

f = 134 × 10⁶N/m²

Therefore,

e = 3000000( 1.078)/ [2(134 × 10⁶) – 3000000]

= 12.2 mm

Other design specifications include:

• Design Pressure = 110% of operating Pressure = 3300 kPa
• Design temperature = Maximum shell temperature = 220^oC
• Corrosion Allowance = 4 mm
  6.       THE QUENCHER

FIGURE 4.2 THE QUENCHER

The quencher is designed as a packed column, so as to increase the contact area between the gas and the water and thus allow for efficient removal of residual carbon from the gas stream.

4.6.1        Design Procedures

The design of a packed column will involve the following steps:

1. Select the type and size of packing.

2. Determine the column height required for the specified separation.

3. Determine the column diameter (capacity), to handle the liquid and vapour flow rates.

4. Select and design the column internal features: packing support, liquid distributor, redistributors.

4.6.2        Types of Packing

The principal requirements of a packing are that it should:

• Provide a large surface area: a high interfacial area between the gas and liquid.
• Have an open structure: low resistance to gas flow.
• Promote uniform liquid distribution on the packing surface.
• Promote uniform vapour gas flow across the column cross-section.

Random packing and structured packing elements are more commonly used in the process industries.

 

DATA GIVEN

Exit Gas stream composition

Moles of exit and inlet gas stream

From material balance, input and exit gas stream composition is calculated.

DATA REQUIRED

Heat Capacity

Amount of water required to achieve the desired temperature change

Tower Diameter

Tower Height

TO OBTAIN MASS OF COOLING WATER REQUIRED:

Heat lost by crude gas= heat gained by cooling water

Entering Crude gas temperature=250^oC

Exiting Crude gas temperature=50^oC

Entering Temperature of cooling water= 25^oC

Exiting temperature of the cooling gas= this is assumed to be 90^oC

Inlet Stream Into Quencher

TABLE 4.7: COMPOSITION OF INLET STREAM INTO QUENCHER

COMPONENT	MOLE (Kmol)	MOLE FRACTION
C	5.16	0.00364
H₂	671.636	0.4743
CO	594.031	0.4195
CO2	117.113	0.0827
N₂	19.754	0.01395
CH₄	1.411	0.000996
H₂S	7.055	0.00498
	1416.16

 

4.6.3        HEAT CAPACITY

The heat capacity data can be obtained from heat capacity tables in available from literature

Taking a reference temperature of 15^oC.

The heat capacity equations are:

C:  Cp = 11.18 + 1.095 × 10^-2T – 4.891 × 10⁵T^-2

H₂: Cp = 

CO: Cp = 

CO₂: Cp = 

N₂: Cp = 

CH₄: Cp = 

H₂S: Cp = 

TABLE 4.8: COMPOSITION OF OUTLET STREAM FROM QUENCHER

COMPONENT	MOLE (Kmol)	MOLE FRACTION
H₂	671.636	0.476
CO	594.031	0.421
CO₂	117.113	0.083
N₂	19.754	0.014
CH₄	1.411	0.001
H₂S	7.055	0.005
	1411

 

Heat lost by the crude gas, ΔH is given by

_net dT

= 1039.525 KJ/Kmol

ΔH_net = (6984.618 × 1416.16) – (1039.525 × 1411) KJ/hr

= 2,340.157 KJ/s

Heat gained by the water = m

= 96,189.91 KJ/Kmol

Thus, Heat gained by water = (m/M x 96189.91KJ/Kmol)

Molecular weight of water = 18 Kg/Kmol

Heat gained by water = 96189.91/18 m = 5343.88 m

Since, heat lost by the crude gas = heat gained by water

Therefore,

5343.88m = 2340.157

m = 0.438 Kg/s

Allowing a minimum of 10 minutes hold-up of water,

Therefore, m = 0.438 Kg/s x 10mins x 60s = 262.748 Kg

4.6.4        COLUMN DIAMETER

D = (4S/π)^1/2

= (4 × 0.2185/3.1429) ^½

= 0.5275 m

 

4.6.5        HEIGHT OF THE COLUMN: Z

This is obtained, from Coulson and Richardson, using the following formula:

Z= H_OGN_OGS.F

= 3.0893 m

7.       INITIAL  H₂S REMOVAL

FIGURE 4.3 THE INITIAL H₂S SCRUBBER

4.7.1        ASSUMPTIONS

• No  NMP  in  Output  gas  stream
• Unit  operates  at  50%  of  the  flooding  gas  mass   velocity
• Flow rate  of  NMP  is  40%  more  than  the  minimum

Given that

Henry’s law constant for NMP = 0.97

Density of gas stream = 31.49kg/m³

Density of NMP = 1020kg/m³

H₂S concentration in the inlet stream = 0.5mol%

NMP has already been selected from the first semester preliminary design as our scrubbing liquid.

An energy balance is not require for this unit because there is very little temperature change between the unit’s inlet and exit streams and hence the unit can be assumed to be operating isothermally.

A packed column using 1 in. Raschig ring packing with a packing factor of 160 have been selected for use for this design. Raschig rings have been selected because they give a larger degree of contact between the gas phase and the liquid phase. It also gives a higher mass transfer area to work with.

Column height = 14m (approx.)

Column diameter = 2.44m (approx.)

 

8.       CARBON MONOXIDE CONVERSION UNIT

FIGURE 4.4 THE CO CATALYTIC CONVERTER

This unit converts carbon monoxide to carbon dioxide, a reaction which leads to the production of more hydrogen. It is required to carry out this reaction over chromium-promoted iron oxide catalyst. The catalyst vessel is a single shell with a dividing plate separating the two catalyst beds which constitute the two stages of conversion. The following reaction takes place in the CO catalytic converter:

CO + H₂O  CO₂ + H₂    ΔH= – 41170.51J

Also provided are the basic data for the CO conversion section of the plant:

1. Space velocity

The space velocity through each catalyst stage be assumed to be 3500 volumes of gas plus steam measured at NTP per volume of catalyst per hour. It should further be assumed that use of this space velocity will allow a 10°C approach to equilibrium to be attained throughout the possible range of catalyst operating temperatures listed below.

2. Equilibrium data for the CO conversion reaction

TABLE 4.9 EQUILIBRIUM DATA FOR CO AT DIFFERENT TEMPERATURES

TEMP/K	600		700		800
KP	3.69 x 10-2		1.11 x 10-4		2.48 x 10-4

In the design of a catalytic converter, the following are the key issues to be considered: reactor type; catalyst type, catalyst volume and mass; the catalytic converter sizing; operating conditions (temperature and pressure) within the reactor; phase type; feed conditions (temperature and concentration).

4.8.1        Reactor Type: Fixed Bed Reactor

The fixed bed reactor is a tubular which is packed with solid granular catalyst on a fixed support. It is termed a fixed bed reactor because the reactor feed, whether liquid or gaseous feed passes through the catalyst bed to contact with the catalyst for chemical reactions, with the catalyst bed not being moved as the fluid flows feed through it. Based on the flow directions of the feedstock some of the hydro – processing reactors can be divided into “axial” reactors and “radial” reactors. In the axial reactor the feed flows along the axial direction for the reactions to take place. Generally the feed flows downwards. In the radial reactor the feed flows in the radial direction for the reactions to take place. The axial reactor is used when the pressure drop across a catalyst bed is small. Or otherwise the radial reactor should be used. Based on the temperature of the cylindrical shell some reactors can be divided into cold wall reactors and hot wall reactors. They are all fixed bed reactors. “Hot wall” reactor is not lined with any heat insulating linings. Because the temperature of the reactor shell is very high, the reactors must be made of alloy steel which can sustain high temperature and is highly resistant to hydrogen corrosion. The advantage is that more volume of such a reactor can be effectively used. A “cold wall” reactor is lined with heat-insulating material. The temperature of the reactor shell is quite low, ordinary low alloy steel can be used for making such reactors. Corrosion caused by feedstock, especially by hydrogen can be avoided.

4.8.2        Catalyst: Chromium – Promoted Iron Oxide

The overall reaction of a heterogeneous gas-solid reaction on a supported catalyst, such as the one involved in this design, is made up of a series of physical steps as well as the chemical reaction. First, there is mass transfer of reactant from the bulk gas phase to the external solid surface, after which there is diffusion from the solid surface to the internal active sites, activation of the adsorbed reactants; next is the most important step of chemical reaction, then desorption of products, internal diffusion of products to the external solid surface, and mass transfer to the bulk gas phase.

TABLE 4.10: CATALYST PROPERTIES OF CHROMIUM PROMOTED IRON CATALYST (Source: Chemical Reactor Design for Process Plants by Howard F. Rase, 1999)

PROPERTY	CHROMIUM – PROMOTED IRON OXIDE
Maximum operating Temperature (K)	890 0F
Tablet size
Bulk Density	70 lb / cu ft
Particle Density	126 lb / cu ft
Cost	20 $ / cu ft
Catalyst poisons	Inorganic salts, boron, oils or phosphorus compounds, liquid H20 is a temporary poison. Sulfur compounds in an amount greater than 50 ppm
Catalyst life	3 years and over depending on care in startup and during operation

The purpose of the chromium promotion is to minimize catalyst sintering by textural promotion.  During the course of the reaction, there are no significant side reactions.

The works of Xue et al confirm that Carbon is formed if the water-gas shift reaction occurs at the reaction stoichiometry and temperature. In order to achieve higher conversions with the use of a catalyst and also avoid coking, the H₂O / CO feed ratio must be higher than that required by the reaction stoichiometry. With this low H₂O / CO ratio, all possible side reactions would likely produce C or CH₄. Formulation of C blocks the catalyst sites causing catalyst deactivation and an increase in the pressure drop across the bed caused by plugging or fouling of the reactor. Formation of CH₄ would consume the hydrogen and alter the product composition which may give rise to potential difficulties in subsequent processes.

Activated Fe₃O₄ is pyrophoric, that is, it ignites spontaneously upon exposure to air, and as a result the catalyst must be re-reduced and stabilized by surface oxidation (using an inert gas with a low concentration of O₂) before being re-used.

4.8.2.1       Catalyst Volume (from space velocity):

The catalyst volume can be obtained from space velocity data given in the problem statement coupled with the simulation figures for the flowrate of gases into the catalytic converter.

FOR 1^ST STAGE:

4.8.2.2       Catalyst Mass (in 1^st stage of converter):

FOR 2^ND STAGE:

4.8.2.3       Catalyst Mass (in 2^nd stage of converter):

4.8.3        CATALYTIC CONVERTER SIZING

In carrying out the sizing of Fixed Bed Reactors, most designs approximate to plug flow. The simplest of such arrangements is adiabatic (Chemical Process Design and Integration by Robin Smith).

A summary of the reactor dimensions is provided in the table that follows.

TABLE 4.11: REACTOR DIMENSIONS OF CATALYTIC CONVERTER

CATALYST BED	LENGTH(M)	DIAMETER(M)	VOLUME(M3)
1st Stage	7.03	3.518	68.34
2nd Stage	12.22	3.518	118.83
REACTOR
1st Stage	8.795	3.518	85.50
2nd Stage	14.61	3.518	142.06
Total	23.405	3.518	227.56

 

 

 

 

9.       FINAL H₂S REMOVAL STAGE

FIGURE 4.5 FINAL H₂S SCRUBBER

The process:

The final H₂S removal stage could be divided into two operations namely:

1. Normal operation
2. Intermittent regeneration operation

4.9.1        NORMAL OPERATION

The converted gas leaving an air-cooled heat exchanger is fed to the unit together with oxygen. Iron oxide in the Lautamasse reacts with hydrogen sulphide in the converted gas to form iron sulphide. The iron sulphide is an intermediate product and it reacts with the oxygen as quickly as it is formed to produce iron oxide and sulphur deposits. This process occurs continuously and is represented by the equation.

2Fe₂O_3(s)  +  6H₂S    2Fe₂S₃   +     6H₂O

2Fe₂S₃      +  3O₂     2Fe2O3 +   6S

4.9.2        REGENERATION OPERATION

When the quantity of sulphur deposit overtime attain a value between 40  55% of the dry weight of the Lautmasse, the normal process will be stopped. Hot oxygen only is then continuously fed to the unit in a parallel in order to oxidize the sulphur deposit to sulphur(iv) oxide gas which can be converted to tetraoxosulphate(vi)acid in order to generate money.

S_(s) +  O_2(g)    SO_2(g)

The Lautamasse becomes ineffective when the regeneration process has been carried out for about 10times.

 

4.9.3        THE EQUIPMENT

The final H₂S removal is to take place in

1. Four vertical vessels
2. A separate outlet from each of the vessel converging SO₂ from regeneration
3. A separate inlet header to each of the column, conveying heated oxygen.
4. A minimum point where the oxygen and the converted gas is mixed.
5. Each column contains five trays each containing Lautamasse
6. The columns are to be identical.

The following was specified

• Height of column      = 18.5 m
• Diameter of column = 2.5 m
• Number of trays       = 5
• Total pressure drop = 35 KN/m²
• Number of columns = 4
• Locking lid type  = Autoclave

Unspecified Parameter are:

1. Plate type and its configuration
2. Mass of iron oxide per tray
3. Regeneration period
4. Regeneration duration
5.   Tray thickness
6. Tray spacing

Choice of tray – Bubble cap

Reasons:

1. It provides maximum contact between the gas and solid unlike the sieve and valve trays.
2. It allows for maximum occupation of the tray space volume by the Lautamasse powder.

1.   The riser hole size  –  5 mm (Coulson and Richardson )
2. Choice of material : Low alloy steel (Ni, Cr, Mo, V)

Properties

Tensile strength   –  550 N/mm²

Design stress       –   240 N/mm²

3.   Determination of number of holes

N = 25001 holes

(e) Determination of the mass of lautamasse and sulphur.

= 10402.6 Kg

When the quantity of sulphur deposit on the tray falls between 40-55% of the weight of the lautamsee the operation has to be stopped and then regeneration takes place.

Assuming the mass of sulphur and the lautamase is the ratio 1:0.4 that is assuming that sulphur is 50% of the total mass mixture.

Mass of sulphur

=3467.5kg

Mass of fe₂0_{3 =}

=6935kg

To calculate the time for regeneration:

4.9.4        ASSUMPTION

The absorption of H₂S in the four columns are 50%, 30%, 20%, 10%. The assumption is based on the fact that absorption decreases along the column. But we want to model the column base on the absorption of the first column.

Time after which regeneration will take place is

=15.8months

16months

After 16 months regeneration process has to be carried out.

Volume and height of sulphur and lautamasse determination.

Mass of sulphur = 3467.6kg

Density of sulphur = 2000kg/m³

Volume of sulphur = 

=

Height = 

= 

= 0.3532m

Lautamasse;

Density of lautamasse = 675kg/m³

Mass of lautamasse = 6935kg

Volume = 

= 10.274m³

Height = 

= 2.0930m

This is a good value since the higher the height of the lautamasse, the lesser the effective tray spacing.

Total height of lautamasse and sulphur

= 2.0930 + 0.3532

= 2.45m

4.9.5        REGENERATION

Sulphur deposit after 16months per column = 3467.5kg  5

= 17337.5kg

Amount in moles= 

= 541.797 moles

S + O  SO₂

541.8 kmol of O₂is required per column.

Since each column is designed assuming 50% of the total H₂S absorbed

(4  541.797kmol)SO₂ = 150% of the actual SO₂ produced

Also, (4  541.797kmol)SO₂ = 150% of the oxygen required

The actual SO₂ produced = 1444.8kmol

The actual O₂ required = 1444.8kmol

If regeneration is assumed to take place within 6hrs,

Flow rate of SO₂ regenerated = 240.8kmol/hr

10.   CARBON DIOXIDE REMOVAL

FIGURE 4.6 CO₂ REMOVAL

Mechanism of Removal: Absorption

Absorption Mechanism: Chemosorption(Chemical Absorption)

Equation of Reaction

K₂CO_3(aq) + H₂O_(l) + CO_2(g)  2KHCO_3(aq)

Basis : 4125.535 kg of Feed / hr

TABLE 4.12: COMPOSITION OF INLET AND OUTLET STREAMS FOR CO₂ REMOVAL

COMPONENT	yin	yout
H2	0.63326	0.9800
CO	0.00060	0.00093
C02	0.35555	0.00247
N2	0.00989	0.0155
CH4	0.00071	0.001104
H2S	5.33E-07	8.34E-07
TOTAL	1.000000	1.000000

ASSUMPTION: Only Carbon dioxide (CO₂)is absorbed and all the other gases act as inert in the 40% K₂CO₃ solution.

Mole Ratio Of Co₂ In The Gas Entering The Column

Mole Ratio Of Co₂  In The Gas Leaving The Column

Plotting the equilibrium curve for CO₂

                     TABLE 4.13: EQUILIBRIUM DATA FOR CO₂ REMOVAL

	X	Y	X	Y
0.025	0.1	0.0017	0.111111	0.001703
0.18	0.2	0.0122	0.25	0.012351
0.6	0.35	0.0408	0.538462	0.042535
1.7	0.5	0.1156	1	0.13071
4.0	0.68	0.2722	2.125	0.374004
7	0.75	0.4762	3	0.909126
10	0.8	0.6803	4	2.127932
11	0.9	0.748	9	2.968254

4.10.1    PACKING MATERIAL: Pall Rings (Plastic)

Pall rings are chosen as the F_p (Packing Factor), small column diameter, high quality with low price, low pressure drop.

The packing characteristics

D_p = 90mm

% = 92%

Summary of design

TABLE 4.14: SUMMARY OF CO₂ ABSORBER DIMENSIONS

COLUMN DIAMETER	1.3m
COLUMN HEIGHT	24.96m
PACKING TYPE	Plastic pall rings
DIAMETER OF PACKING	90mm
VOIDAGE	92%
SPECIFIC AREA	85m2/m3
PACKING FACTOR	17.07 ft-1
QTY OF SOLVENT REQUIRED	852.537 kmol/hr

 

11.   HEAT EXCHANGER DESIGN

4.11.1    THE PROPOSED DESIGN

TABLE 4.15: SUMMARY OF PROPOSED HEAT EXCHANGER DESIGN

EXCHANGER	FIRST	SECOND
Tube Layout	Split ring, floating head, 1 shell pass, and 1 tube pass of4.83 m length, 30 mm O.D., 26 mm I.D., and triangular pitch. Pitch-to-diameter ratio of 1.25	Split ring, floating head, 1 shell pass, and 4 tube passes of6.00 m length, 20 mm O.D., 16 mm I.D., and triangular pitch. Pitch-to-diameter ratio of 1.25
Total number of Tubes/Number of Tubes per pass	96/96	1308/327
Heat transfer area (based on outside diameter)m2	43	493
Shell Internal diameter (mm)	484.5	1064.4
Baffle spacing (mm)	97	212.88
Baffle cut	25%
Tube-side coefficient (W/m2 °C), clean.	252.97	102.42
Overall coefficient, estimated (W/m2 °C).	120.00	60.00
Overall coefficient required (W/m2°C).	149.85	67.81
Dirt/Fouling factors  Tube-side/Shell-side(m2 °C /W)	0.0002 /0.0003	0.0002 /0.0003
Pressure drops, including drop over nozzles (estimated)  Tube-side/Shell-side (kPa)	253/452	73/969

 

4.11.2    OPTIMISATION

There is scope for optimizing the design by reducing the number of tubes, as the pressure drops are well within specification and the overall coefficient is well above that needed. However, the method used for estimating the coefficient and pressure drop on the shell side(Kern’s method) is not so accurate, so keeping to this design will give some margin of safety.

4.11.3    VISCOSITY CORRECTION FACTOR

The viscosity correction factor  was neglected when calculating the heat transfer coefficients and pressure drops. This is reasonable for gases as it has a relatively low viscosity.

Usually, 

For such a small factor, the decision to neglect it was justified. Applying the correction would decrease the estimated heat transfer coefficient, which should not be a problem as the area required for heat exchange is well taken care of. It would also give a slight increase in the estimate.

4.11.4    OXYGEN STREAM PRE-HEATER.

4.11.4.1   FLUID ALLOCATION

Placing the higher temperature stream on the tube side will reduce the overall cost of design. Therefore, the saturated stream is placed in the tubes.

4.11.4.2   DUTY OF PRE HEATER

Let us assume that the amount of saturated stream used for pre-heating of the O₂ stream gives off ‘ONLY’ its Latent Heat to the O₂ – stream

Q= 877.04 ×10³ KJ/hr

= 243.622 ×10³ W

 

MASS OF SATURATED STREAM REQUIRED.

m_S =  0.1414 Kg/sec

 

 

MEAN TEMPERATURE DIFFERENCE (ΔTm)

ΔT_m= 0.986(114.823 0c)

= 133.215 ⁰c

4.11.4.3   ESTIMATION OF U (Overall Coefficient)

An estimation of 200W/m²C for the overall Heat Transfer Coefficient is chosen as suggested by literature.

4.11.4.4   PROVISIONAL AREA.

A  = 10.76m²

4.11.4.5   TUBE DIMENSION SELECTION

• Outer Diameter – 20mm
• Inner Diameter – 16mm
• Length of Tube – 6.10m
• Tube Effective Length- 6.0m

Small diameter tubes are chosen because they give more compact & cheaper exchangers. Use of longer tubes will reduce shell diameter thus lowering cost exchanger.

Heat Transfer Area of 1 tube = πd_oL

= π(20/1000)(6)

= 0.37699m²

4.11.4.6   NUMBER OF TUBES REQUIRED

Number of tubes = 28.542

=28 tubes

4.11.4.7   TUBE ARRANGEMENT SELECTION.

A Triangular Tube Arrangement pattern is chosen as it gives higher heat transfer rates as compared to others. The recommended tube pitch is 1.25 times the Tube Outside Diameter(do). Tube Bundle Diameter is therefore

= 170mm

4.11.4.8   SHELL DIAMETER DETERMINATION

Using a Split Ring Floating Head unit, the Diameter clearance between the Shell and Tube Bundle is 50mm

Therefore,

Shell Diameter = 170 +50 = 220mm

4.11.4.9   TUBE SIDE COEFFICIENT

A Horizontal Condenser with condensation in the tubes is the usual arrangement for pre-heaters &vaporizers using condensing steam as the heating medium. It is customary to design purposes. For air- free steam, a coefficient of 8000W/m² is used. Therefore,

h_I= 8000W/m²

4.11.4.10                   MASS VELOCITY IN SHELL (G_S)

G_S = 134.917Kg/m²s

4.11.4.11                   LINEAR VELOCITY (Us)

U_S = 3.20m/sec

 

4.11.4.12                   SHELL SIDE EQUIVALENT DIAMETER

For a Triangular pitch Arrangement

d_e = 1.10/d_o[pt²– 0.917d_o²]

= 14.2mm

 

4.11.4.13                   AVERAGE PHYSICAL PROPERTIES OF SHELL FLUID

• Viscosity (µ) = 2.61 ×10^-5 Pas
• Thermal Conductivity (K) = 3.448 × 10^-2 W/mK
• Heat Capacity (Cp) = 0.9818 ×10³ J/Kg ⁰C

4.11.4.14                   REYNOLD’s NUMBER ON SHELL SIDE

Re = 73402.99987

4.11.4.15                   PRANDTL’s NUMBER ON SHELL SIDE

Pr = 0.74318

4.11.4.16                   HEAT TRANSFER COEFFICIENT (SHELL SIDE)

h_s = 317.324W/m^2 0c

4.11.4.17                   OVERALL COEFFICIENT

U= 322.691 W/m^2 0c

4.11.4.18                   PRESSURE DROP CALCULATIONS

On the Shell Side

Re= 73402.99987

J_F= 3.6 × 10^-2

ΔP_s = 4(3.6×10^-2)(220/14.2)(6000/220)(42.07×3.20²)

= 26.2112 KN/m² which is very acceptable

On the Tube Side:

The pressure drop on the condensing side or tube side is difficult to predict as 2 phases are present & the vapour mass velocity is changing throughout the tube side.

4.11.5    AIR-COOLED EXCHANGERS

Air-cooled exchangers should be considered when cooling water is in short supply or expensive. They can also be competitive with water-cooled units even when water is plentiful. Frank (1978) suggests that in moderate climates air-cooling will usually be the best choice for minimum process temperatures above 65°C, and water-cooling for minimum process temperatures below 50°C. Between these temperatures a detailed economic analysis would be necessary to decide the best coolant. Air-cooled exchangers are used for cooling and condensing. Air-cooled exchangers consist of banks of finned tubes over which air is blown or drawn by fans mounted below or above the tubes (forced or induced draft). Typical units are shown in Figure 2 below. Air-cooled exchangers are packaged units, and would normally be selected and specified in consultation with the manufacturers. The equation for finned tubes given below can be used to make an approximate estimate of the area required for a given duty (Sinnott, 2005).

FIGURE 4.7 AIR-COOLED EXCHANGERS

4.11.5.1   FINNED TUBES

Fins are used to increase the effective surface area of heat-exchanger tubing. Many different types of fin have been developed, but the plain transverse fin shown in Figure 2 is the most commonly used type for process heat exchangers.

FIGURE 4.8: FINNED TUBE

Finned tubes are used when the heat-transfer coefficient on the outside of the tube is appreciably lower than that on the inside; as in heat transfer from a liquid to a gas, such as in air-cooled heat exchangers.

The fin surface area will not be as effective as the bare tube surface, as the heat has to be conducted along the fin. This is allowed for in design by the use of a fin effectiveness, or fin efficiency, factor.

(Sinnott, 2005)

     (4.11.1)

LMTD = 47.28°C

      (4.11.2)

Prandtl Number

      (4.11.3)

Thus, the Prandtl number for air at 216.4 °C = 0.69

Heat Duty Required

      (4.11.4)

Hence, the heat duty Q of the exchanger = 1.013 KJ/kg °C

Overall Heat Transfer Coefficient for the Design

      (4.11.5)

Trials with various values of U

TABLE 4.16: HEAT TRANSFER COEFFICIENTS WITH CORRESPONDING AREA AND FACE VELOCITY VALUES

S/N	U (KJ/hr m2°C)	A (m2)	vair (m/s)
1	6.127 × 104	965.56	3.00
2	8.169 × 104	724.167	4.00
3	1.225 × 105	482	6.00

In addition,

      (4.11.7)

A_face = face area (m²)

TABLE 4.17: CORRESPONDING FACE AREA AND FACE VELOCITY VALUES

vface (m/s)	Aface (m2)
3	1135.96
4	851.97
6	567.91

Substituting W_air in (4.11.7) into (4.11.4),

      (4.11.8)

ρ_air = density of air, 1.1762 kg/m³

So far, for heat exchanger parameters,

U = 1.2254 × 10⁵ KJ/hr m² °C

A = 482.78 m²

v_face = 6 m/s

A_face = 567.91 m²

4.11.5.2   FAN COVERAGE

= 293.96 m²

For the design also, we choose to use two fans rather than one, as this takes care of cooling over a bit of a longer distance of fan stages, and provides insurance in case of failure of one fan or the other.

4.11.5.3   FAN DIAMETER

With the two identical fans to be used the diameter is simply computed with the diameter, thus;

Diameter of fan = 13.45 m

General Properties of the Fans to be used as per industrial standards

Fan speeds v_fan = 70 m/s

V-belt drive of about 30 horsepower

Forced draft fan system will be used as it uses less power in blowing cool air volume

4.11.5.4   FIN AREA AND TUBE AREA ANALYSES

         (4.11.10)

     (4.11.11)

      (4.11.12)

D_fin = Diameter of fin = 12.7 mm

D_e = External diameter of the tube = 41.2 mm

D_i = Internal diameter of the tube = 37.2 mm

W = Width of tube = 4.00 m

L_fin = Length of fin = 68.2 mm

L = Length of tube = 15 m

A_HT = Total heat transfer area per length of tube = 0.01644 m²/m

Number of tubes per row of each metre in a tube bank = n^× = 

With the chosen pitch, P = 47.625 mm; n^× = 20.997 ≈ 21 tubes per row

Number of rows n, chosen = 4; and given that W = 4.00 m,

Total number of tubes, N = n^× × W × n = 336 tubes.

12.   PUMP SIZING

Because liquids are incompressible, Equation 1 may be used to calculate the work required to pump a liquid. The kinetic energy term is small compared to the other terms and is neglected. Therefore, Equation 4.11.1 reduces to

g∆z + ∆p/ρ+ W+ E = 0,  defining points 1 and 2,we have,

g (z₂ – z₁,) + (p₂ – p₁,)/ρ + W + E = 0       (4.12.1)

The friction pressure loss term is split into two parts, one for the suction side of the pump, Es, and, the other for the discharge side of the pump, E_D. Thus, Equation 1 becomes, after solving for W.

W = g (z₁ – z₂) + (p₁ – p₂)/ρ – (Es + E_D)       (4.12.2)

Frequently, we must make a preliminary estimate of the pump work. Manufacturers do not stock all pumps and other expensive machinery because of the cost of carrying an inventory. The machinery is manufactured on receipt of an order from a customer. Manufacturing some process machinery, may take six months or longer. To save time in implementing a project will require ordering equipment having long delivery times before completing a detailed design. Also, the management of a firm will require an estimate of the cost of a project to prepare a budget or a proposal for a customer. Once the work is estimated, the pump shaft power,

P_p = mW/ ɳ_P          (4.12.3)                                                                

where,  ɳ_p – the pump efficiency, includes both the hydraulic and mechanical frictional losses. Below outlines a calculation procedure for calculating an approximate pump size.

4.12.1    Approximate Pump-Sizing Calculation Procedure

1. Define the flow system, i.e., locate points 1 and 2. The pressures p₁ and p₂ will be known at these points.

2. Locate the process equipment according to the rules-of-thumb which is zero above ground level for pumps.

3. Estimate z ₁ and z₂.

4. Estimate the frictional pressure losses Es and E_D according to the table 2 given below.

5. Calculate the pump work from Equation (4.12.2)

6. Calculate the pump shaft horsepower using Equation (4.12.3) and the pump efficiencies

given in the  excel spreadsheet.

7. Calculate the electric-motor horsepower using the motor efficiency given in

the spreadsheet.

8. Select a standard electric-motor horsepower using the spread sheet to obtain approximately

a 10% safety factor.

 

 

TABLE 4.18: APPROXIMATE FRICTIONAL PRESSURE DROP ACCROSS PROCESS EQUIPMENT

4.12.2    THE METERING RAM PUMP

The pumped fluid contacts only the inside surface of the tubing. There are no other valves, O-rings, seals or packing to worry about in a peristaltic pump. Therefore, the only compatibility to worry about in a peristaltic pump is the tubes for the fluid being pumped. Of all the design parameters, the compatibility is the first one that need to be considered.

The tube is made elastomeric in order to maintain the circular cross section of cycles of squeezing in the pump. This requirement eliminates a variety of non-elastomeric polymers that have compatibility with a wide range of chemicals, such as PTFE, polyolefins, PVDF etc. from consideration as material for pump tubing. Elastomers, both natural and synthetic, have common characteristics including elastic recovery after stress, flexibility, extrusion resistance, and relative impermeability to gases and liquids. Each class of elastomer has its own unique properties and performance that can be modified by the inclusion of other ingredients.

Reasons why tube is made elastomeric and not plastic are because elastomer provide :

1. Very good dimensional stability over a broad temperature range

2. High compression set resistance

3. Excellent extrusion resistance at high temperatures and pressures

4. Retention of sealing force during pressure and temperature cycling (a major problem with many plastics)

5. Elastomers are long-chain polymers connected by cross-links that impart strength, resilience and elasticity, and these cross-links are very stable to heat and high pressure. Plastics are also long-chain polymers, but they are not connected by chemical cross-links. Plastics acquire their strength when the chains orient with one another and become crystalline in regions. These regions may deform or melt under high pressure and temperature.

There are over 20 classes of elastomers. Of all these classes, the most ideal elastomer for this work is EPDM (ethylene propylene diene Monomer) + polypropylene (as in Santoprene).

Reasons why EPDM (ethylene propylene diene Monomer) + polypropylene (as in Santoprene) is the most ideal :

1. It has a wide service temperature range

2. EPM elastomers have excellent resistance to ozone, water and steam, alkalis and acids, salt solutions and oxygenated solvents

3. Very good balance of compression set, flex resistance, physical strength, low temperature flexibility, weather resistance properties, and resistance to acid condensates

4. Outstanding resistance to high heat; excellent resistance to oil, gasoline, hydraulic fluids and hydrocarbon solvents; very good impermeability to gases and vapor; very good resistance to weather, oxygen, ozone, and sunlight; good flame retardant.

5. EPDM + polypropylene have the best fatigue resistance compare to other elastomers

4.12.2.1   OCCLUSION

The minimum gap between the roller and the housing determines the maximum squeeze applied on the tubing. The amount of squeeze applied to the tubing affects pumping performance and the tube life – more squeezing decreases the tubing life dramatically, while less squeezing decreases the pumping efficiency, especially in high pressure pumping. Therefore, this amount of squeeze becomes an important design parameter.

The term “occlusion” is used to measure the amount of squeeze. It is either expressed as a percentage of twice the wall thickness, or as an absolute amount of the wall that is squeezed.

Mathematically;

Where;

For this design work , the wall thickness of the tube of the peristaltic pump(t) is 4mm and the minimum gap between the roller and the housing(g) is 1.6mm so the percentage occlusion is calculated using the above formular

Occlusion is one of the key peristaltic pump design parameter because it affects not only the efficiency of the fluid movement in the pump but the life span of the tube. The higher the occlusion, the more efficiently the pump moves fluid and the lower the occlusion, the less efficiently the pump moves fluid. The occlusion not only alters the efficiency of the pump but also the lifespan of the tubing. By increasing the occlusion, you are squeezing the tubing with greater force, which will wear out the tubing material faster. Prepare replacement tubing are provided because of the high occlusion , and because a replacement tubing may be needed is one of the reasons why EPM elastomer tubing is  the choice of tubing because it is inexpensive.

CHAPTER FIVE

5           MECHANICAL DESIGN

1.       THE GASIFIER

5.1.1        DESIGN PRESSURE

5.1.2        DESIGN TEMPERATURE

5.1.3        MATERIAL OF CONSTRUCTION

The wall of the gasifier is made of Austenitic stainless steel (Type 316) as in the storage tanks.

Lining Material (Refractory block):Refractory liners are used on the working face of entrained-flow slagging gasifiers that react coal, petroleum coke, heavy fuel oil or other feedstock with oxygen and water. The refractory liners protect the gasifier shell from elevated temperatures, corrosive slags, and thermal cycling during gasification. Refractory failure is primarily by two means which are corrosive dissolution and spalling. High chrome gasifier materials have evolved as the material of choice to line the hot face of gasifiers, yet the performance of these materials does need to improve by coating it with substances like magnesium oxide, aluminium oxide and/or chromium oxide (Bennet, 2004).

Therefore,

The refractory liner for the heavy fuel oil gasifier is: Chromium oxide coated with zirconium oxide and aluminium oxide (Cr₂O₃: 90%, Al₂O₃: 3%, ZrO₂: 7% in weight %)

5.1.4        THICKNESS OF THE REFRACTORY BRICK

Assuming that the ratio of the thickness of the vessel to the lining material is taken as 8:1, the thickness of the refractory lining is therefore = 90.6/8 mm = 11.325 mm

5.1.5        MINIMUM WALL THICKNESS

The minimum thickness required to resist the internal temperature of the gasifier is calculated thus,

5.1.6        MINIMUM HEAD THICKNESS

An Ellipsoidal head is chosen as the closure for this vessel because it is relatively cheaper than the torispherical head and it can withstand the design pressure of the vessel (Coulson and Richardson, 1999). The minimum thickness required is obtained thus,

Thickness of the head = 92.2 mm

5.1.7        VESSEL SUPPORT

Skirt Support is chosen because it is suitable for tall and vertical columns as it does not impose concentrated load on the vessel.

2.       THE WASTE HEAT BOILER

With the heat exchanger geometry defined, the mechanical design calculations must be done to ensure that the heat exchanger design is valid for the design pressure and conditions. The typical calculations are:

• Calculation of shell wall thickness.
• Calculation of nozzle wall thickness.
• Calculation of tube plate thickness.

SHELL WALL THICKNESS DETERMINATION

P_i= 3000 kPa

D_i = 1.078 m

At the shell side mean temperature of 192^oC,

f = 134 × 10⁶N/m²

Therefore,

e = 3000000( 1.078)/ [2(134 × 10⁶) – 3000000]

= 12.2 mm

Other design specifications include:

• Design Pressure = 110% of operating Pressure = 3300 kPa
• Design temperature = Maximum shell temperature = 220^oC
• Corrosion Allowance = 4 mm
  3.       THE QUENCHER

5.3.1        MATERIALS FOR CONSTRUCTION:

Carbon Steel is advised because, of its operability at not so high-pressure, because of its hardness and malleability and because it is most popular and therefore easy to obtain.

5.3.2        PACKING:

Interlox saddle packing was chosen because they are more efficient due to greater surface and improved hydrodynamics. They are made with a variety of holes and protrusions to enlarge the specific surface area. Because of their open structure and large specific surface area, mass transfer efficiency is high when proper distribution of liquid and gas over the cross section can be maintained. It also important to note that when ceramic construction of packing is suitable, saddles are the preferred packing.

5.3.3        DESIGN PRESSURE

Assuming 8% above operating pressure

= 1.08 x 4241.325 KN/m²

= 4580 .631 KN/m²

(4580.631-101.325) gauge

= 4479.306 kN/m² gauge

5.3.4        DESIGN TEMPERATURE

This is obtained as,

= Highest Temperature + allowance for uncertainty

Assuming an uncertainty of 80^oC

Therefore, design temperature= (250+ 80) ^oC

=330^oC

5.3.5        TENSILE STRENGTH

At 25^oC, the tensile strength of Carbon =210N/mm²

At 500^oC, the tensile strength of Carbon is 450 N/mm²

At 330^oC, the tensile strength of Carbon will be,

TABLE 5.1: TENSILE STRENGTH VALUES AT DIFFERENT TEMPERATURES

TEMPERATURE, OC	TENSILE STRENGTH, N/MM2
25	210
330	X
500	450

Interpolating we have,

(330 – 25) / (500 – 25) = (x – 200) / (450 – 210)

x=154+200

=  354N/mm²

5.3.6        DESIGN STRESS:

At 300^oC, the design strength of Carbon = 85N/mm²

At 350^oC, the design strength of Carbon is 80 N/mm²

At 330^oC, the design strength of Carbon will be,

(COULSON AND RICHARDSON, VOL 6)

TABLE 5.2: DESIGN STRESS VALUES AT DIFFERENT TEMPERATURES

TEMPERATURE, OC	DESIGN STRESS, N/MM2
25	85
330	X
500	80

Interpolating we have,

(330 – 300) / (350 – 300) = (x – 85) / (80 – 85)

X = τ = 82 N/mm²

Nominal design stress = 1.5 × = 1.5 x 82

= 123 N/mm²

5.3.7        CHOICE OF CLOSURE

Minimum wall thickness required for a cylindrical shell

e= Pi Di / 2f – Pi

Given that,

P_i = 4241.325 KN/m²

D_i = Internal diameter = 0.5275 m

therefore,

E = (4241.325 × 0.5275)

2 × 82000 – 4241.325

= 14 mm

Corrosion allowance = 2 mm

Minimum thickness of wall required = 14 mm + 2 mm = 16 mm

Choice of closure = Hemispherical head

Optimum thickness ratio for a hemispherical head to the thickness of the vessel = 0.6 (Coulson and Richardson, 2005)

Thickness of the hemispherical head = 16 x 0.6 mm = 9.6 mm

Type of Vessel Support = Skirt Support

 

4.       THE INITIAL H₂S ABSORBER

5.4.1        Shell Thickness

Thickness of shell, 

where,

Inner diameter of vessel (D) = 1.3 m

Working pressure = 1.013 ×10⁶ N/m²

Design pressure (P) = 1.1× 1.013 x 10⁶ N/m²

Permissible Stress (f) = 115 × 10⁶ N/m²

Joint Efficiency (J) = 0.85

Corrosion allowance (c) = 2mm

hence,

So outer diameter of shell

1.                Axial Stress Due to Pressure

Axial stress due to pressure, 

2.                Stress due to Dead Load

1.                 Compressive Stress due to weight of shell up to a distance X

 

2.                Compressive stress due to weight of insulation at height X

Insulator used is asbestos

Thickness of insulation, = 100mm

Diameter of insulation, 

Density of insulation = 575 kg/m³

Mean diameter of vessel = 

For large diameter column,

3.                 Compressive stress due to liquid in column up to height X

4.                Effect of attachments

1. Packing weight
2. Head weight
3. Ladder

Density of packing (Plastic pall rings) = 53 kg/m³

Packing weight = 

= 690.11 X N

Head weight (approximately) = 35000 N

Weight of ladder = 1600 X N

5.4.2        SUPPORT FOR ABSORBER

Material used is structural steel (IS 800)

Skirt support is used

Inner diameter of the vessel = D_i = 4.69m

Outer diameter of the vessel = D_o = 4.70m

Height of the vessel = H = 10.054m

Density of carbon steel = = 7700kg/m³

Density of packing =  = 53 kg/m³

5.       THE CO CONVERTER

The catalytic converter primarily functions to provide surface area required for the exothermic reaction involved in the water-gas shift reaction and also to house the catalyst required for the reaction. The materials that make entry into the converter include steam and flue gases which exit the initial H₂S removal stage and basically include hydrogen, carbon monoxide. Hydrogen sulphide, methane and nitrogen. The major aspects that will be considered in the mechanical design include:

• Operating and design temperatures and pressures
• Vessel dimensions and orientation
• Construction materials
• Type of vessel heads

5.5.1        Operating and Design temperatures and pressures

The operating temperature of the catalytic converter: 380 ⁰C

The operating pressure of the catalytic converter: 1200 kPa

The design temperature of the catalytic converter: 480 ⁰C (catalyst)

The design pressure of the catalytic converter: 1.1  1200 kPa = 1320 kPa

The maximum operating with respect to catalyst degeneration:  476.67 ⁰C

5.5.2        Vessel Dimensions and Orientation

The major dimensions have earlier been calculated in the chemical engineering design section of this unit; the diameter being 3.518m and total height of 23.41m. Minimum wall thickness – the thickness required to ensure that any vessel is sufficiently rigid to withstand its own weight, and any incidental loads; and for a diameter of about 3.5 m, this thickness is about 16 mm. Also, a 3mm corrosion allowance is added. The “corrosion allowance” is the additional thickness of metal added to allow for material lost by corrosion and erosion, or scaling.

5.5.3        Construction Materials

Construction materials for pressure vessels are usually plain carbon steels; low and high alloy steels, other alloys, clad plate, and reinforced plastics. Because the temperature of the reactor shell is very high, the reactors must be made of alloy steel which can sustain high temperature and is highly resistant to hydrogen corrosion. A “cold wall” reactor is lined with heat-insulating material. The temperature of the reactor shell is quite low, ordinary low alloy steel can be used for making such reactors.

Alloy steels contain one or more alloying agents to improve mechanical and corrosion-resistant properties over those of carbon steel.

Nickel increases toughness and improves low-temperature properties and corrosion resistance. Chromium and Silicon improve hardness, abrasion resistance, corrosion resistance, and resistance to oxidation.

Molybdenum provides strength at elevated temperatures.

 

 

5.5.4        TYPE OF VESSEL HEAD

The vessel head indicates the closure. Ellipsoidal heads will be made use of. The ellipsoidal head thickness can be calculated using the following formula:

where,  P_i – is the internal pressure; = 1200 kPa =1.2 Nmm^-2

D_i – the internal diameter; = 3.518 m

 J – the joint factor; which is equal to 1 as a result of the absence of joints

 f– the design stress at the specified temperature, equal to 180 N/mm²

Others include:

5.5.5        Openings required

Two openings are required:

An entry opening at the vessel top

An exit opening at the vessel bottom

Removable baffles for catalyst loading and discharge

5.5.6        Heating and Cooling Jackets or Coils.

Since its been established that a cold wall reactor will be used for the wall of the catalytic converter, implying an adiabatic arrangement to be ensured by incorporating on the reactor body an insulating material.

5.5.7        Internal fittings specification

The reactor will also consist of a reactor cylindrical shell, inlet diffuser, vapor-liquid distributor, scale removing basket, catalyst supporting plate, catalyst pipe, quench hydrogen box and a re-distributor, outlet accumulator, catalyst removal port and quench hydrogen pipe.

Also needed are Skirt supports of about 5 mm thickness;

FIGURE 5.1 SKIRT-SUPPORT

Base rings and anchor bolts for the supports.

FIGURE 5.2 DOUBLE PLATE WITH GUSSET

6.       THE FINAL H₂S ABSORBER

Design stress                  240N/mm₂

Construction material ;  low alloy steel

Tray thickness;              12mm

Design temperature;      30 – 45^oC

5.6.1        Design pressure:

This should be 10% greater than operating pressure

=  

5.6.2        Wall thickness

e = 

e = wall thickness

Pi= internal design pressure

Di= internal chamber

= design stress

e = 

= 12.9 mm

Using corrosion allowance of 2mm

e= 14.9mm

5.6.3        Support

Skirttype is recommended

7.       THE CO₂ ABSORBER

Material for the absorber shell is Carbon steel

5.7.1        Shell Thickness

Thickness of shell, 

where,

Inner diameter of vessel (D) = 1.3 m

Working pressure = 1.013 ×10⁶ N/m²

Design pressure (P) = 1.1× 1.013 x 10⁶ N/m²

Permissible Stress (f) = 115 × 10⁶ N/m²

Joint Efficiency (J) = 0.85

Corrosion allowance (c) = 2mm

Hence,

So outer diameter of shell

5.7.2        Axial Stress Due to Pressure

Axial stress due to pressure , 

5.7.3        Stress due to Dead Load

5.7.3.1       Compressive Stress due to weight of shell up to a distance X

    

5.7.3.2       Compressive stress due to weight of insulation at height X

Insulator used is asbestos

Thickness of insulation, = 100mm

Diameter of insulation, 

Density of insulation = 575 kg/m³

Mean diameter of vessel = 

For large diameter column,

5.7.3.3       Compressive stress due to liquid in column up to height X

5.7.4        Effect of attachments

4. Packing weight
5. Head weight
6. Ladder

Density of packing (Plastic pall rings) = 53 kg/m³

Packing weight = 

= 690.11 X N

Head weight (approximately) = 35000 N

Weight of ladder = 1600 X N

5.7.5        SUPPORT FOR ABSORBER

Material used is structural steel (IS 800)

Skirt support is used

Inner diameter of the vessel = D_i = 4.69m

Outer diameter of the vessel = D_o = 4.70m

Height of the vessel = H = 10.054m

Density of carbon steel = = 7700kg/m³

Density of packing =  = 53 kg/m³

8.       THE HEAT EXCHANGERS

With the heat exchanger geometry defined, the mechanical design calculations must be done to ensure that the heat exchanger design is valid for the design pressure and conditions. The typical calculations are:

• Calculation of shell wall thickness.
• Calculation of nozzle wall thickness.
• Calculation of inner tube wall thickness.
• Calculation of expansion joint dimensions (to compensate for shell and tube side differential expansion due to temperatures differences.
• Calculation of tube plate thickness.

5.8.1        FIRST EXCHANGER

5.8.1.1       Shell Wall Thickness  (Corrosion Allowance = 2 mm):

 (Sinott, 2005)

Where f is the design stress, and P_i is the internal pressure.

Design pressure; take as 10 per cent above operating pressure,

= 31.9 bar

= 3.19 N/mm²

Typical design stress f = 540 N/mm²

e = = 1.70; with J = 0.85 for shell and tube exchangers

Adding corrosion allowance, 1.70 + 2 = 3.70, say 4.00 mm

5.8.1.2       Nozzle wall thickness:

Tube has been designed to provide entrance for nozzle = (0.2 shell diameter)

Thus nozzle diameter =

Nominal thickness at this condition is 0.5 in (=12.7mm) (Shilling et al., 1997)

5.8.1.3       Inner tube wall thickness:

With the tube dimensions selected according to standard,

Thickness = 

 

 

 

5.8.1.4       Calculation of tube plate thickness:

The minimum plate thickness enough to withstand the bending and shearing force is gotten from literature as:

Minimum plate thickness = 0.75

5.8.1.5       OTHER SPECIFICATIONS

Number of baffles    Hence  

Baffle clearances

The edge distance between the outer-tube limit (OTL) and the baffle diameter has to be sufficient to prevent tube breakthrough due to vibration.

Baffle clearance = 11mm (for shell ID m) (Shilling et al., 1997)

Distance from tube O.D to centre of partition

Distance = 9.5mm (for ID)

Design temperature

Built to withstand 10% above highest stream temperature = (373.4 °C) × 1.1 ≈ 411 °C

Design pressure

Take as 10% above operating pressure = (30-1) × 1.1 = 39.1 bars

 

SECOND EXCHANGER

Shell wall thickness e (Corrosion Allowance = 2 mm):

(Sinott, 2005)

Design pressure; take as 10 per cent above operating pressure,

= 26.4 bar

= 2.64 N/mm²

Typical design stress f = 540 N/mm²

e = = 3.07; recall J = 0.85

Adding corrosion allowance, 3.07 + 2 = 5.07, say 6.00 mm

Nozzle wall thickness:

Tube has been designed to provide entrance for nozzle = (0.2 shell diameter)

Thus nozzle diameter=

Nominal thickness at this condition is 0.432in (=10.97mm)

Inner tube wall thickness:

Thickness=

Calculation of tube plate thickness:

Minimum plate thickness=0.75

OTHER SPECIFICATIONS

Number of baffles    Hence  

Baffle clearances

Baffle clearance=13mm (for shell IDm) (Shilling et al., 1997)

Distance from tube O.D to centre of partition

Distance = 9.5mm (for IDm)

 

Design temperature

Built to withstand 10% above highest stream temperature = (380.9 °C) × 1.1 ≈ 420 °C

Design pressure

Take as 10% above operating pressure = (25-1) × 1.1 = 26.4 bars

FIGURE 5.3: PHYSICAL LAYOUT OF A 2-TUBE PASS SHELL AND TUBE EXCHANGER

CHAPTER SIX

6           PLANT HAZARD AND OPERABILITY STUDY

1.       GASIFIER

GASIFIER INTENTION: to cause partial oxidation of the heavy oil feedstock.

TABLE 6.1:HAZOP ANALYSIS FOR GASIFIER

GUIDE WORD	DEVIATION	CAUSES	CONSEQUENCES	ACTION
NO	Flow (steam & gas)	Control valve stuck closed	Delays the reaction	Fit low temperature alarm
MORE	Flow (Steam/Oxygen)	Failure of ratio flow or failure of control valve	High reactor temperature Disaster	1)Alarms  2) Automatic shutdown
AS WELL AS	Composition	Refractory particles from reactor	Possible blockage of boiler tubes	Install filters up-stream of boiler

 

 

2.       WASTE HEAT BOILER

TABLE 6.2: HAZOP ANALYSIS FOR WASTE HEAT BOILER

GUIDE WORD	DEVIATION	CAUSES	CONSEQUENCES	ACTION
NONE	No tube side fluid flow	Failure of inlet tube side  valve to open	Process fluid temperature is not heated accordingly	Install Temperature indicator before and after the process fluid line  Install TAH
MORE	More tube side fluid flow	Failure of inlet tube side valve to close	Output of Process fluid temperature too high	Install Temperature indicator before and after process fluid line  Install TAL