Process Guidelines
Short Description
Process Guidelines...
Description
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN
Rev.: A
PROCESS DESIGN GUIDELINES
Date: 08/02/08 Page 1 of 87
WORK IMPROVEMENT PLAN
PROCESS DESIGN GUIDELINES
WIP-SIPS-PCS-001
08/02/08
A
First Issue
RM PCS
GP PCS
GP ENG
Date
Revision
Description of Revision
Prepared by
Checked by
Approved by
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN PROCESS DESIGN GUIDELINES
Rev.: A
Date: 08/02/08 Page 2 of 87
REVISION RECORDING
Date
Revision
Description of Revision
Prepared by
Checked by
Approved by
08/02/08
A
First Issue
RM PCS
GP PCS
GP ENG
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN PROCESS DESIGN GUIDELINES
Rev.: A
Date: 08/02/08 Page 3 of 87
TABLE OF CONTENTS 1.
INTRODUCTION.......................................................................................................................4
2.
RESPONSIBILITY ....................................................................................................................4
3.
SCOPE......................................................................................................................................4
4.
PHYSICAL PROPERTIES ........................................................................................................5
5.
FORMULA ................................................................................................................................6
6.
INSULATION SPECIFICATION FOR PIPING..........................................................................9
7.
MATERIALS OF CONSTRUCTION .......................................................................................10
8.
SELECTION OF THERMODYNAMIC MODEL IN HYSYS.....................................................11
9.
PIG LAUNCHER AND RECEIVER.........................................................................................14
10.
PUMP SELECTION AND SYSTEM DESIGN .........................................................................16
11.
VESSEL SELECTION AND SIZING.......................................................................................25
12.
HEAT EXCHANGERS ............................................................................................................28
13.
COLUMNS & TOWERS..........................................................................................................34
14.
COMPRESSORS AND VACUUM EQUIPMENT ....................................................................36
15.
SAFETY SYSTEM & PSV DESIGN........................................................................................38
16.
PIPE SIZING ...........................................................................................................................43
17.
SELECTION OF VALVES ......................................................................................................48
18.
FLARE SYSTEM ....................................................................................................................57
19.
HEAT TRANSFER FLUID (HTF) SYSTEM DESIGN .............................................................71
20.
COOLING WATER SYSTEM DESIGN...................................................................................72
21.
REFRIGERATION SYSTEMS ................................................................................................74
22.
CHILLED WATER SYSTEM...................................................................................................75
23.
CHILLED BRINE SYSTEM.....................................................................................................76
24.
DM WATER SYSTEM.............................................................................................................77
25.
STEAM SYSTEM ....................................................................................................................78
26.
PLANT & INSTRUMENT AIR SYSTEM .................................................................................80
27.
NITROGEN GENERATION SYSTEM ....................................................................................82
28.
INCINERATOR SYSTEM .......................................................................................................83
29.
EFFLUENT TREATMENT SYSTEM.......................................................................................84
30.
PROCESS CONTROLS .........................................................................................................86
31.
REFERENCE ..........................................................................................................................87
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN PROCESS DESIGN GUIDELINES
Rev.: A
Date: 08/02/08 Page 4 of 87
1.
INTRODUCTION This guideline generally outlines the methods for designing the process systems.
2.
RESPONSIBILITY The design engineer shall be responsible for carrying out the Process system design. The respective lead engineer shall counter check the correctness of calculations and design.
3.
SCOPE This guideline covers the following subjects in detail: 1) Formulas 2) Material of construction 3) Storage tank selection and sizing 4) Vessel selection and sizing 5) Pump selection and sizing 6) Line sizing and selection 7) Control valve selection and sizing 8) Compressor selection 9) Safety valve relief load calculation and selection 10) Raw water / service water / potable water system design 11) Cooling water system design 12) Chilled water / chilled brine system design 13) DM water system design 14) Steam & condensate system design 15) Service and Instrument air system design 16) Nitrogen generation / storage system design 17) Oxygen storage and system design 18) Diesel / Fuel oil / LSHS storage and handling system design 19) Flare system design and selection 20) Chemicals storage and handling system design 21) Process control
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN
Rev.: A
PROCESS DESIGN GUIDELINES
Date: 08/02/08 Page 5 of 87
4.
PHYSICAL PROPERTIES
Property Heat Capacity
Units
Water
Organic Liquids
Steam
Air
Organic Vapors
KJ/kg °C
4.2
1.0 - 2.5
2.0
1.0
2.0 - 4.0
Btu/lb °F
1.0
0.239 - 0.598
0.479
0.239
0.479 - 0.958
kg/m³
1000
700 - 1500
1.29@STP
lb/ft³
62.29
43.6 - 94.4
0.08@STP
KJ/kg
1200 - 2100
200 - 1000
Btu/lb
516 - 903
86 - 430
W/m °C
0.55 - 0.70
0.10 - 0.20
0.025 - 0.070
0.025 - 0.05
0.02 - 0.06
Btu/h ft °F
0.32 - 0.40
0.057 - 0.116
0.0144 - 0.040
0.014 - 0.029
0.116 - 0.35
cP
1.8 @ 0°C
**See Below
0.01 - 0.03
0.02 - 0.05
0.01 - 0.03
10-1000
1.0
0.7
0.7 - 0.8
Density
Latent Heat
Thermal Cond. Viscosity
0.57 @ 50°C 0.28 @ 100°C 0.14 @ 200°C Prandtl Number
1 - 15
** Viscosities of organic liquids vary widely with temperature Liquid density varies with temperature by:
ρ L ⋅ ∝ ⋅(TC − T ) 0.3 Gas density can be calculated by:
ρG =
MW × P Z ⋅ R ⋅T
Boiling Point of Water as a Function of Pressure:
(
Tbp = P × 10 9
)
0.25
;
Tbp in °C ; P in MPa
Density of Metals Metal
Density (kg/m³)
Aluminum
3500
Carbon Steel
7800
Galvanized Iron
6000
Stainless Steel
8000
Titanium
4000
FRP
2000
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN PROCESS DESIGN GUIDELINES
Rev.: A
Date: 08/02/08 Page 6 of 87
5.
FORMULA
5.1
VOLUME AND SURFACE AREA Volume and surface area for different shapes and sections are: Section
Volume
Surface area
Sphere
πD³/ 6
πD²
Hemi-head
πD³ / 12
πD²/ 2
S.E.head
πD³ / 24
1.084D²
Ellipsoidal head
πD² l / 6
2πR² + (πl / e) ln ((1+ e)/(1 – e))
100 – 60% F& D head
0.08467D³
0.9286D²
F & D head
2πR³ K / 3
πR² ( 1+ l2/ R² (2 – l/R))
Cone
πD² l / 12
πDl / 2cosά
Truncated cone
(π l (D² + D d + d²) ) / 12
π((D + d)/2) sqrt (l²+((D – d)/2)²)
30° Truncated cone
0.227(D³ – d³)
1.57(D²– d²)
Cylinder
πD² l / 4
πDl
2
. Where l = Height of cone, depth of head or length of cylinder ά = one- half apex angle of cone D = Large diameter of cone / diameter of head or cylinder R = Radius r = Knuckle radius of F & D head L = Crown radius of F & D head h = Partial depth of horizontal cylinder K,C = Coefficients d = Small diameter of truncated cone V = Volume
K=
L 2r ⎞ ⎛ L ⎞⎛ L − ⎜ − 1⎟⎜ + 1 − ⎟ R R⎠ ⎝ R ⎠⎝ R
e = 1−
l2 R2
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN PROCESS DESIGN GUIDELINES
Rev.: A
Date: 08/02/08 Page 7 of 87
5.2
PARTIAL VOLUME FOR HORIZONTAL VESSELS Volume of cylinder (V1)
⎞ ⎛ ∝° V1 = L ⋅ R 2 ⋅ ⎜ − sin(∝ °) × cos(∝ °) ⎟ ⎠ ⎝ 57.3 h cos(∝ °) = 1 R ∝ ° = Degrees{ A cos(∝)} Volume of Head (V2) (both heads)
V2 = 2 × 0.215 × h × (3R − h ) Total Volume (V) V = V1 + V2 5.3
PARTIAL VOLUME FOR VERTICAL VESSELS Volume of cylinder (V1)
V1 =
π 4
D2h
Volume of Head (V2) (one head)
V2 = 0.215 × h × (3R − D ) Total Volume (V) V = V1 + V2 5.4
WETTED AREA FOR HORIZONTAL VESSELS Wetted area of cylinder (A1)
A1 = R ⋅ 2 ⋅ ( Ar cos(1 − h / R) ) × L Wetted area of heads (A2) (heads are assumed to be hemispherical) (two heads)
⎛ R−H⎞ A2 = 2πR 2 ⋅ ⎜1 − ⎟ R ⎠ ⎝ Total wetted area (A) A = A1 + A2
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN PROCESS DESIGN GUIDELINES
Rev.: A
Date: 08/02/08 Page 8 of 87
5.5
WETTED AREA FOR VERTICAL VESSELS Wetted area of cylinder (A1) A1 = π ⋅ D ⋅ h Wetted area of heads (A2) (heads are assumed to be hemispherical) (one head) A2 = 2 ⋅ π ⋅ R 2 Total wetted area (A) A = A1 + A2 Where D = Diameter of the vessel (m) R = Radius of the vessel = D/2 (m) L = Length of the vessel (m) h = Height of liquid from vessel bottom (m) A1 = Wetted area of cylindrical portion (m²) A2 = Wetted area of heads (m²) A = Total wetted area of the vessel (m²) V1 = Partial volume of cylindrical portion (m²) V2 = Partial volume of heads (m²) V = Total liquid volume of the vessel (m²)
5.6
TWO PHASE DENSITY AND VISCOSITY The density and viscosity of mixed phase fluid is found by the following method:
λ = Ql / (Ql + Q g ) ρ h = ρ l λ + ρ g (1 − λ ) µ h = µ l λ + µ g (1 − λ ) Where Q l g h ρ µ
= Volumetric flow rate (m³/h) = Liquid = Gas or vapor = Mixed phase or homogenous phase = Density (kg/m³) = Viscosity (cP)
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN
Rev.: A
PROCESS DESIGN GUIDELINES
Date: 08/02/08 Page 9 of 87
6.
INSULATION SPECIFICATION FOR PIPING
6.1
INSULATION THICKNESS FOR HOT SERVICE The maximum fluid temperature (°C) is shown as a function of the insulation thickness and pipe size.
Insulation Thickness (mm)
0
25
40
50
65
80
Nominal Pipe Size
90
100
115
125
140
150
165
180
190
200
Maximum Fluid Temperature (°C)
1/2”
7
68
230
456
723
760
1”
7
61
205
409
653
760
1 ½”
7
48
150
302
491
707
760
2”
7
45
134
269
439
635
760
2 ½”
7
42
122
244
398
579
760
3”
7
40
111
221
361
526
710
760
4”
7
37
100
195
319
465
629
760
6”
7
35
86
164
265
386
524
676
760
8”
7
34
79
147
236
343
465
599
745
760
10”
7
33
75
136
216
312
422
544
676
760
12”
7
32
72
129
203
292
393
506
629
760
14”
7
32
70
125
196
281
379
487
605
731
760
16”
7
32
68
121
188
268
360
462
573
692
760
18”
7
31
67
117
181
258
345
442
548
661
760
20’
7
31
66
114
176
249
333
426
527
635
750
760
Over 24” and flat surface
7
31
64
110
167
236
314
400
494
594
701
760
To obtain the correct insulation thickness, read across from the correct pipe diameter to a tabulated process temperature that is greater or equal to the actual temperatures. Read up for the insulation thickness. For example a 16” diameter pipe operating at 190°C will require 80 mm of insulation. 6.2
INSULATION MATERIAL • Up to 343°C (650°F), magnesia is most used • From 871°C to 1037°C (1600°F – 1900°F) a mixture of asbestos and diatomaceous earth is used • Ceramic refractories are used for higher temperatures • Cryogenic equipment -129°C (-200°F) employs insulation with fine pores in which air is trapped
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN PROCESS DESIGN GUIDELINES
Rev.: A
Date: 08/02/08
Page 10 of 87
7.
MATERIALS OF CONSTRUCTION Material
Carbon Steel
Stainless Steel
254 SMO (Avesta)
Titanium
Advantage Low cost, easy to fabricate, abundant, most common material. Resists most alkaline environments well. Relatively low cost, still easy to fabricate. Resist a wider variety of environments than carbon steel. Available is many different types. Moderate cost, still easy to fabricate. Resistance is better over a wider range of concentrations and temperatures compared to stainless steel. Very good resistance to chlorides (widely used in seawater applications). Strength allows it to be fabricated at smaller thicknesses.
Disadvantage Very poor resistance to acids and stronger alkaline streams. More brittle than other materials, especially at low temperatures. No resistance to chlorides and resistance decreases significantly at higher temperatures. Little resistance to chlorides, and resistance at higher temperatures could be improved. While the material is moderately expensive, fabrication is difficult. Much of cost will be in welding labor.
Pd stabilized Titanium
Superior resistance to chlorides, even at higher temperatures. Is often used on sea water application where Titanium's resistance may not be acceptable.
Very expensive material and fabrication is again difficult and expensive.
Nickel
Very good resistance to high temperature caustic streams.
Moderate to high expense. Difficult to weld.
Hastelloy Alloy
Very wide range to choose from. Some have been specifically developed for acid services where other materials have failed.
Fairly expensive alloys. Their use must be justified. Most are easy to weld.
Graphite
One of the few materials capable of withstanding weak HCl streams.
Brittle, very expensive, and very difficult to fabricate. Some stream components have been known to diffusion through some types of graphites.
Tantalum
Superior resistance to very harsh services where no other material is acceptable.
Extremely expensive, must be absolutely necessary.
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN
Rev.: A
PROCESS DESIGN GUIDELINES
Date: 08/02/08
Page 11 of 87
8.
SELECTION OF THERMODYNAMIC MODEL IN HYSYS Application
Type
Model
Applicable Range
Remarks
OIL & GAS Reservoir Systems
Equations of state for high pressure hydrocarbon applications
Platform Separation
Equations of state for high pressure hydrocarbon applications
Transportation of oil and gas by pipeline
Equations of state for high pressure hydrocarbon applications
Peng-Robinson (PENG-ROB)
All Pressures
All Temperatures
Should not be used if polar components such as Alcohols are present.
Redlich-KwongSoave (RKSOAVE)
All Pressures
All Temperatures
Should not be used if polar components such as Alcohols are present.
Peng-Robinson (PENG-ROB)
All Pressures
All Temperatures
Should not be used if polar components such as Alcohols are present
Redlich-KwongSoave (RKSOAVE)
All Pressures
All Temperatures
Should not be used if polar components such as Alcohols are present
Peng-Robinson (PENG-ROB)
All Pressures
All Temperatures
Should not be used if polar components such as Alcohols are present
Redlich-KwongSoave (RKSOAVE)
All Pressures
All Temperatures
Should not be used if polar components such as Alcohols are present
Braun K-10 (BK10)
Suited for vacuum and low pressure applications
Applicable range is 133 o 800 K. Can be used up to o 1100 K
Gives fast and acceptable answers. Should be use as a first attempt only. This model should not be used for systems containing Hydrogen.
REFINERY
Low Pressure applications (up to several atm) eg: Vacuum tower & crude tower
Petroleum Correlation Models
Chao-Seader (CHAO-SEA)
Grayson
Chao-Seader (CHAO-SEA)
Medium Pressure applications (up to several tens of atm) Coker Main Fractionator FCC Main Fractionator
< 140 atm
< 210 atm
< 140 atm
o
200-533 K
o
200-700 K
o
200-533 K
Petroleum Correlation Models
Equation Of State Models
o
This model should not be used for systems containing Hydrogen. If light ends dominate, Equation of State models can be used. This model should be used for systems containing Hydrogen. If light ends dominate, Equation of State models can be used. This model should not be used for systems containing Hydrogen. If light ends dominate, Equation of State models can be used. This model should be used for systems containing Hydrogen. If light ends dominate, Equation of State models can be used.
Grayson
< 210 atm
200-700 K
Peng-Robinson (PENG-ROB)
All Pressures
All Temperatures
Should not be used if polar components such as Alcohols are present
Redlich-KwongSoave (RKSOAVE)
All Pressures
All Temperatures
Should not be used if polar components such as Alcohols are present
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN
Rev.: A
PROCESS DESIGN GUIDELINES
Date: 08/02/08
Page 12 of 87
Hydrogen Rich Applications eg: Reformer, Hydrofiner, Hydrotreaters & HydroDesulfurisers
Lube Oil Unit, De-Asphalting Unit
Petroleum Correlation Models
Equation Of State Models
Equation Of State Models
o
This model should be used for systems containing Hydrogen. If light ends dominate, Equation of State models can be used
Grayson
< 210 atm
200-700 K
Peng-Robinson (PENG-ROB)
All Pressures
All Temperatures
Should not be used if polar components such as Alcohols are present
Redlich-KwongSoave (RKSOAVE)
All Pressures
All Temperatures
Should not be used if polar components such as Alcohols are present
Peng-Robinson (PENG-ROB)
All Pressures
All Temperatures
Should not be used if polar components such as Alcohols are present
Redlich-KwongSoave (RKSOAVE)
All Pressures
All Temperatures
Should not be used if polar components such as Alcohols are present
Peng-Robinson (PENG-ROB)
All Pressures
All Temperatures
Should not be used if polar components such as Alcohols are present
Redlich-KwongSoave (RKSOAVE)
All Pressures
All Temperatures
Should not be used if polar components such as Alcohols are present
Peng-Robinson (PENG-ROB)
All Pressures
All Temperatures
Should not be used if polar components such as Alcohols are present
Redlich-KwongSoave (RKSOAVE)
All Pressures
All Temperatures
Should not be used if polar components such as Alcohols are present
Peng Robinson MHV2 equation of state (PRMHV2)
Accurate up to 150 bar
All Temperatures
Can be used for mixture of polar and non polar components
Peng Robinson Wong-Sandler Equation Of State (PRWS)
Accurate up to 150 bar
All Temperatures
Can be used for mixture of polar and non polar components
Predictive Soave-RedlichKwong (PSRK)
All Pressures
All Temperatures
Can be used for mixture of polar and non polar components
Redlich Kwong MHV2 equation of state (RKSMHV2)
Accurate up to 150 bar
All Temperatures
Can be used for mixture of polar and non polar components
Redlich Kwong Wong-Sandler Equation Of State (RKSWS)
Accurate up to 150 bar
All Temperatures
Can be used for mixture of polar and non polar components
Chao-Seader (CHAO-SEA)
< 140 atm
200-533 K
Grayson
< 210 atm
200-700 K
Peng-Robinson (PENG-ROB)
All Pressures
All Temperatures
GAS PROCESSING Hydrocarbon separations Demethanizer C3splitter
Cryogenic gas processing
Gas Dehydration with glycols
Equation Of State Models
Equation Of State Models
Flexible and Predictive Equation of State Model
PETROCHEMICALS Ethylene Plant Main Fractionator
Petroleum Correlation Models
Ethylene Plant Light Hydrocarbon
Equation Of State Models
o
o
This model should not be used for systems containing Hydrogen. This model should be used for systems containing Hydrogen. Should not be used if polar components such as Alcohols are present
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN PROCESS DESIGN GUIDELINES
Rev.: A
Date: 08/02/08
Page 13 of 87 Separation Train
Ethylene Plant Quench Tower
Aromatics (eg: BTX Extraction)
Substituted hydrocarbons, VCM Plant & Acrylo Nitrile Plant
Ether Production eg: MTBE, ETBE, TAME
Equation Of State Models
Liquid Activity Coefficients
Equation Of State Models
Liquid Activity Coefficients(very sensitive to parameters)
Equation Of State Models
Ethyl Benzene and Styrene Plants Liquid Activity Coefficients
Redlich-KwongSoave (RKSOAVE)
All Pressures
All Temperatures
Should not be used if polar components such as Alcohols are present
Peng-Robinson (PENG-ROB)
All Pressures
All Temperatures
Should not be used if polar components such as Alcohols are present
Redlich-KwongSoave (RKSOAVE)
All Pressures
All Temperatures
Should not be used if polar components such as Alcohols are present
NRTL
Low pressures up to 10 atm.
No component should be close to its critical temperature.
UNIFAC
Low pressures up to 10 atm.
No component should be close to its critical temperature.
UNIQUAC
Low pressures up to 10 atm.
No component should be close to its critical temperature.
Peng-Robinson (PENG-ROB)
All Pressures
All Temperatures
Should not be used if polar components such as Alcohols are present
Redlich-KwongSoave (RKSOAVE)
All Pressures
All Temperatures
Should not be used if polar components such as Alcohols are present
NRTL
Low pressures up to 10 atm.
No component should be close to its critical temperature.
UNIFAC
Low pressures up to 10 atm.
No component should be close to its critical temperature.
UNIQUAC
Low pressures up to 10 atm.
No component should be close to its critical temperature.
Peng-Robinson (PENG-ROB)
All Pressures
All Temperatures
Should not be used if polar components such as Alcohols are present
Redlich-KwongSoave (RKSOAVE)
All Pressures
All Temperatures
Should not be used if polar components such as Alcohols are present
NRTL
Low pressures up to 10 atm.
No component should be close to its critical temperature.
UNIFAC
Low pressures up to 10 atm.
No component should be close to its critical temperature.
UNIQUAC
Low pressures up to 10 atm.
No component should be close to its critical temperature.
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN PROCESS DESIGN GUIDELINES
Rev.: A
Date: 08/02/08
Page 14 of 87
9.
PIG LAUNCHER AND RECEIVER LAUNCHER
RECEIVER Kiker line
BL
Kiker line
AL
AR
BR
Typical diameters of Major Barrel and Pipework: Pipeline diameter
Bypass line
Kicker line
Balance line
Drain line
Major barrel
4”
3”
2”
2”
2”
6”
6”
4”
2”
2”
2”
8”
8”
4” – 6”
4”
2”
2”
10”
10”
6”
4”
2”
2”
12”
12”
6” – 8”
4”
2”
2”
16”
14”
6” – 10”
4”
2”
2”
16”
16”
8” – 12”
6”
4”
4”
18”
18”
10” – 12”
8”
4”
4”
20”
20”
10” – 16”
8”
4”
4”
24”
24”
12” – 18”
8”
4”
4”
28”
28”
16” – 20”
10”
4”
4”
32”
30”
16” – 24”
10”
4”
4”
36”
32”
16” – 24”
10”
4”
4”
36”
36”
18” – 28”
12”
4”
4”
40”
38”
20” – 28”
12”
4”
4”
42”
40”
20” – 32”
12”
4”
4”
44”
42”
20” – 36”
16”
4”
4”
46”
48”
24” – 36”
18”
4”
4”
52”
56”
32” – 40”
20”
4”
4”
60”
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN
Rev.: A
PROCESS DESIGN GUIDELINES
Date: 08/02/08
Page 15 of 87
Barrel lengths for Intelligent Pigs: Approx. minimum barrel length (m) (Notes 1 & 2)
Approx. maximum tool length (m) (Note 2)
Approx. maximum tool weight (kg) (Note 3)
AL
BL
AR
BR
4”
2.8
60
2.8
0.5
2.8
2.8
6”
2.8
90
2.8
1.5
2.8
2.8
8”
3.9
170
4.1
1.5
3.9
3.9
10”
4.3
300
4.3
1.5
4.3
4.3
12”
4.3
365
4.3
1.5
4.3
4.3
14”
4.8
380
4.8
1.5
4.8
4.8
16”
5.1
700
5.1
1.5
5.1
5.1
18”
5.1
810
5.1
1.5
5.1
5.1
20”
5.1
840
5.1
1.5
5.1
5.1
24”
5.7
1600
5.7
1.5
5.7
5.7
28”
5.8
2000
5.8
1.5
5.8
5.8
30”
6.0
2000
6.0
1.5
6.0
6.0
32”
6.6
2270
6.6
1.5
6.6
6.6
36”
6.6
3560
6.6
1.5
5.3
6.6
38”
6.6
3600
6.6
1.5
5.5
6.6
40”
6.6
4090
6.6
1.5
5.5
6.6
42”
6.6
4550
6.6
1.5
6.4
6.6
48”
6.6
Note 4
6.6
1.5
6.6
6.6
56”
6.6
Note 4
6.6
1.5
6.6
6.6
Pipeline diameter
Launcher
Receiver
NOTES: 1) Refer the figure above for details 2) The lengths are extreme figures. To be checked with supplier for accurate dimensions 3) The weight is indicative for the pig only, excluding the weight of lifting/loading trolley 4) To be checked with supplier
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN PROCESS DESIGN GUIDELINES
Rev.: A
Date: 08/02/08
Page 16 of 87
10.
PUMP SELECTION AND SYSTEM DESIGN
10.1
PUMP SELECTION Service Very high head at low flow Metering small flows (< 1 m³/h) Viscous fluid (> 1000 cP) Non-Newtonian fluid Suspension of crystals or fragile solids Entrained gas (> 2 vol%)
•
• • •
10.2
Type of Pump Reciprocating Reciprocating (plunger or diaphragm) Rotary gear or screw Screw Rotary lobe Rotary or diaphragm
Centrifugal pumps: Single stage for 3.4 - 1134 m³/h (15 - 5000 GPM) & 152 m maximum head Multi Stage for 4.6 - 2500 m³/h (20 - 11,000 GPM) & 1675 m maximum head Efficiencies of 45% at 23 m³/h (100 GPM), 70% at 113 m³/h (500 GPM), 80% at 2270 m³/h (10,000 GPM). Axial pumps can be used for flows of 4.6 - 22680 m³/h (20 - 100,000 GPM) Expect heads up to 12 m and efficiencies of about 65-85% Rotary pumps can be used for flows of 0.23 - 1134 m³/h (1 - 5000 GPM) Expect heads up to 15,200 m (50,000 ft) and efficiencies of about 50 - 80% Reciprocating pumps can be used for 2.3 - 22680 m³/h (10 - 100,000 GPM) Expect heads up to 300,000 m (1,000,000 ft). Efficiencies:70% at 7.46 kW (10 hp), 85% at 37.3 kW (50 hp) and 90% at 373 kW (500 hp)
PUMP PERFORMANCE WITH IMPELLER AND SPEED CHANGE Diameter change
Q2 = Q1 ×
D2 D1
⎛D H 2 = H 1 × ⎜⎜ 2 ⎝ D1
⎞ ⎟⎟ ⎠
2
⎛D BHP2 = BHP1 × ⎜⎜ 2 ⎝ D1
⎞ ⎟⎟ ⎠
3
⎞ ⎟⎟ ⎠
3
Speed change
Q2 = Q1 ×
N2 N1
⎛N H 2 = H 1 × ⎜⎜ 2 ⎝ N1
⎞ ⎟⎟ ⎠
2
⎛N BHP2 = BHP1 × ⎜⎜ 2 ⎝ N1
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN PROCESS DESIGN GUIDELINES
Rev.: A
Date: 08/02/08
Page 17 of 87
Dia & Speed change ⎛D N ⎞ Q2 = Q1 × ⎜⎜ 2 × 2 ⎟⎟ ⎝ D1 N 1 ⎠
⎛D N ⎞ H 2 = H 1 × ⎜⎜ 2 × 2 ⎟⎟ D ⎝ 1 N1 ⎠
2
⎛D N BHP2 = BHP1 × ⎜⎜ 2 × 2 ⎝ D1 N 1 10.3
⎞ ⎟⎟ ⎠
3
EFFECT OF VISCOSITY ON PUMP PERFORMANCE • • • •
Alternate for centrifugal pump (rotary pump) to be considered when fluid viscosity is above 220 cSt (or 2500 SSU). Small pumps become impractical above 220 cSt. Generally centrifugal pumps are limited to about 4000 SSU max. viscosity Correction factors to head, capacity and efficiency should be applied when viscosity is above 70 SSU. Viscosity correction calculation: Viscous capacity (gpm) Qvis = Qwater* Cq Viscous head (ft) Hvis = Hwater* Ch Viscous efficiency Evis = Ewater* Ce Correction factors (Cq,Ch,Ce)
Cq=0.95-0.9 @200-1000 cSt Ch=0.85-0.95 @200-1000 cSt Ce=0.55-0.75 @200-1000 cSt
BHP @normal service = (Q*H*SG)/(2.31*1750*E) BHP @viscous service = (Qvis*Hvis*SG)/(2.31*1750*Evis) 10.4
EFFECT OF VAPOR OR GASES 1) Most centrifugal pumps can handle up to 3 vol. of vapor per 100 vol. of liquid with the max. limit at 7 - 8%. 2) If vapor is suspected in the pump suction area, specify any of the following: - Self-priming type pump - Increase suction pressure - Specify low speed pump - Specify that pump casing should be capable of accepting an oversize impeller to counter the resulting loss in head and capacity.
10.5
MINIMUM FLOW BYPASS 1) Single stage pumps, bypass = 15 - 25 % Multistage pumps, bypass = 25 - 35 % Worthington suggests 30 gpm per 100 HP 2) The temperature rise in a turbine regenerative pump is much greater than in a conventional centrifugal pump, since the HP input increases as the flow rate thro the turbine pump is decreased, where as the HP input to a centrifugal pump decreases as flow is reduced. Consequently a turbine pump is always provided with bypass and also a relief valve 3) The increase in liquid temperature at low flow rates may be calculated by assuming that all the HP shown on the pump curve at the desired capacity is being converted to heat, except that which is used to deliver the small capacity (by the following formulae).
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN PROCESS DESIGN GUIDELINES
Rev.: A
Date: 08/02/08
Page 18 of 87
∆T = At shut-off
∆T =
(BHP
fromcurve
− 0.000252 × Q × H × Sg ) × 42.4 Q × 8.33 × Sg × Cp
(BHP
at ⋅ shutoff
)× 42.4
W × Cp
Where ∆T = Temp diff (°F), Q = gpm, H = head (ft), Sg = Specific gravity; Cp = Specific heat capacity (Btu/lb°F), W = weight of fluid in pump (lb) 4) Minimum flow orifice sizing can be done by the following formulae:
Q = 29.8 × K × d 2 × ⎛⎜ ∆P ⎞⎟ ⎝ Sg ⎠ Where Q = gpm, K = orifice disc coef =0.65, d = orifice dia (inch), ∆P = pres. drop (psi), Sg = Specific Gravity 10.6
MINIMUM NPSH® FOR CENTRIFUGAL PUMPS Pump capacity (m³/h) Up to 11.4 11.4 – 22.7 22.7 – 45.5 45.5 - 91 91 - 159
NPSHr (m) 1.5 1.8 2.1 – 2.4 2.7 – 3.0 3.0 – 3.7
Pump capacity (m³/h) 159 - 227 227 - 455 454 - 568 > 568
NPSHr (m) 3.7 – 4.3 4.3 – 6.1 6.1 – 7.3 7.6
Required NPSH can be reduced as follows: - Use a double suction pump - Use a slower speed - Use smaller pumps in parallel - Use an oversize pump - Use impeller with inducer
10.7
PUMP CENTERLINE ELEVATION FOR CENTRIFUGAL PUMPS Pump capacity (m³/h) Up to 45.5 45.5 – 227
10.8
Centerline Elevation (m) 0.76 0.91
Pump capacity (m³/h) 227 - 2270 2270 - 4545
Centerline Elevation (m) 1.1 1.4
Pump capacity (m³/h) 45.5 – 114 114 - 227 227 - 1136
Efficiency (%) 50 – 75% 60 – 80% 70 – 82%
CENTRIFUGAL PUMP EFFICIENCY Pump capacity (m³/h) Up to 11.4 11.4 – 22.7 22.7 – 45.5
Efficiency (%) 20 – 40% 30 – 50% 35 – 65%
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN PROCESS DESIGN GUIDELINES
Rev.: A
Date: 08/02/08
Page 19 of 87
10.9
CENTRIFUGAL PUMP MOTOR RATING Shaft Power (BkW) Up to 22.4 22.4 – 74.6 > 74.6
Motor Rating 1.25 x BkW 1.15 x BkW 1.10 x BkW
10.10 STEAM TURBINE DRIVEN CENTRIFUGALPUMPS Shaft Power (BkW) 37.3 74.6 149.2
Efficiency (%) 33 37 45
Shaft Power (BkW) 373 746
Efficiency (%) 58 64
10.11 PRESSURE DROP FOR EQUIPMENT AND PIPIING Equipment / Piping Item Exchangers / Air coolers / Double pipes Pump suction screen Rotary & turbine flow meters Flow orifice
Pump suction Pump discharge 0 - 57 m³/h 57 - 159 m³/h > 159 m³/h
Pressure Drop (bar) 0.7 0.07 0.5 0.2
Pressure drop 0.23 - 0.68 bar/km
Pressure drop
Carbon Steel Pipe 5.7 – 22.6 bar/km 3.4 – 15.8 bar/km 2 – 9.1 bar/km
Alloy Steel Pipe 13.6 – 33.9 bar/km 13.6 – 33.9 bar/km 4.5 – 15.8 bar/km
Higher velocity is considered for bigger pipes and for higher operating pressures. 10.12 MULTIPLYING FACTOR FOR EQUIVALENT LENGTH Pipe dia up to 3" 4" 6" 8" and above
30 1.9 2.2 2.7 3.4
Approx. straight length (m) 60 1.6 1.8 2.1 2.4
150 1.2 1.3 1.4 1.6
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN PROCESS DESIGN GUIDELINES
Rev.: A
Date: 08/02/08
Page 20 of 87
10.13 PUMP SEALS In order to prevent leakage of pump fluid through the clearance between the pump shaft and casing, two types of sealing device are generally used in pumps namely; (1) Gland Packing; (2) Mechanical Seal Gland Packing: This is a simple packing by a sealing device which is a rope like material wound between the pump shaft and casing. Frequent changeover of packing is required to prevent leakage. This is used for domestic and water applications where leakage is not a concern. Mechanical Seals Single mechanical seals are generally used to arrest leakage, without frequent maintenance. Double mechanical seals are sometimes preferred in toxic, flammable and corrosive services. 1) Packed seal (Mechanical seal) Packed seal does not require an external flushing liquid if the pumped liquid has lubricating properties at the seal conditions. An external flushing system is required for the following: - For vacuum service to prevent air intake. - For abrasive containing fluids. - For handling volatile fluids which vaporize at operating conditions and present a fire hazard. - For pumping toxic or corrosive liquids. 2) Seal piping - Seal cooler duty can be estimated by for a seal fluid flow of 0.5 m³/h per seal. - For dead end seals the seal fluid flow can be 0.23 m³/h per seal - Seal fluid pressure can be 0.7 – 1.72 bar above packing box pressure. Turbine or pump seal, gland, stuffing box cooling water rate shall be 1-2 m³/h. Refer API-610 to understand more about the type of flushing plans for each service. 10.14 CONTROL VALVE PRESSURE DROP FOR CENTRIFUGAL PUMPS 1) Minimum pressure drop for a control valve is 0.7 bar @design flow rate. 2) Generally CV loss shall be taken as 25% of the total system friction loss or 10% of operating pressure 3) More detailed approach for CV pressure drop is:
dPv = A × dPf + B × dPs Where: dPv dPf dPs A, B Qd Qm
= Control valve pressure drop = Frictional pressure drop of system = Pressure differential due to static head of system = Multiplication factors (as given below) = Design / rated flow rate = Maximum process flow rate
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN
Rev.: A
PROCESS DESIGN GUIDELINES
Date: 08/02/08
Page 21 of 87
Pump Over design Factor (Qd / Qm)
A
B
0-5% 10% 15% 20% 25%
0.294 0.468 0.658 0.875 0.950
0.059 0.094 0.132 0.175 0.200
10.15 PUMP SYSTEM CURVE
Hso Hd dPv
dPo
dPf
dPs
Qm in
Qm
Qn
Qd
Where: dPv = Control valve pressure drop at maximum flow Qm dPo = Control valve pressure drop at full open condition at design flow Qd
10.16 PUMP CALCULATION Pb
H3 Pa
dPf2 dPf1 H1
PS
H2 Pd
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN PROCESS DESIGN GUIDELINES
Rev.: A
Date: 08/02/08
Page 22 of 87
Where: Pa Pb Pv Ps Pd Sg dPf1 dPf2 H1 H2 H3
= Suction vessel minimum operating pressure (barg) = Discharge vessel maximum operating pressure (barg) = Vapor pressure of fluid at maximum operating temperature (bara) = Pump suction pressure (barg) = Pump discharge pressure (barg) = Specific Gravity of the fluid at operating conditions = Suction piping frictional loss (bar) = Discharge piping frictional loss (bar) = Suction vessel minimum liquid level (LSLL) elevation (m) (bar) = Pump suction nozzle elevation (m) (bar) = Discharge vessel maximum liquid level (LSHH) elevation (m) (bar)
To convert elevation in meters to pressure unit = H / 10.2 * Sg (in bar) Suction pressure (Ps) Discharge pressure (Pd) Pump differential pressure (DP) Pump head (h)
= Pa + H1 – H2 – dPf1 = Pb + H3 – H2 + dPf2 = Pd – Ps = DP *10.2 / Sg
Available NPSH (NPSHa)
= (Pa + 14.7) – Pv – dPf1 + (H1 - H2) (bara) = NPSHa *10.2 / Sg (m)
(barg) (barg) (bar) (m)
General considerations: 1) RVP is lower than the True vapor pressure (TVP or Absolute vapor pressure) of hydrocarbon liquids. Correction chart is used to find out TVP from RVP at operating temperature. TVP is used for pump calculations 2) Calculated NPSHa - 1 m (less 1 m) should be given in the process datasheet. The difference between vendor NPSHr and datasheet NPSHa should not be less than 1 m. 3) In absence of data pump suction elevation can be considered as 0.6 m for smaller pumps and 1.0 m for larger pumps (> 200 m³/h). 4) In absence of data take 0.1 bar line loss for suction piping or 0.05 bar for smaller pumps. 5) Spare single stage pumps in service over 260°C and spare multistage pumps in service over 200°C must be kept warm for quick start-up, by providing a bypass around check valve. Needle valve to be provided for the bypass. 6) Pumping capacity over design: Column overhead pumps 20% Product and transfer pumps 10% Design pressure of pump = Design pressure of suction vessel (PSV set) + Max. suction static head + Max. pump differential pressure Pump Power calculation Power estimates for pumping liquids BkW = { Flow (m³/h) x Head (m) x Sp.Gr. } / {367 x Efficiency} BHP = { Flow (US gpm) x Head (ft) x Sp. Gr. } / {3960 x Efficiency} **Efficiency expressed as a fraction in these relations An equation developed for efficiency based on the GPSA Engineering Data Book is:
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN PROCESS DESIGN GUIDELINES
Rev.: A
Date: 08/02/08
Page 23 of 87
Efficiency = 80-0.2855F+0.000378FG-0.000000238FG^2+0.000539F^2-0.000000639(F^2)G + 0.0000000004(F^2)(G^2)
Where; Efficiency is in fraction form, F = developed head (ft), G = flow (gpm) Ranges of applicability are F = 15 - 91 m and G = 23 – 230 m³/h 10.17 PUMP OVERDESIGN FACTOR Service Product Reflux Circulating reflux
Over design Factor 10% 15% 5%
Service Sour water drain Intermittent Utilities
Over design Factor 25% 0% 10%
10.18 RECIPROCATING PUMP a) Simplified formulas for sizing dampers (for pump speeds up to 100 rpm) Pump type Simplex, single-acting Simplex, double-acting Duplex, double-acting Triplex, single & double-acting
Damper volume, US Gallons V = 5 x discharge rate in gpm / rpm V = 2.5 x discharge rate in gpm / rpm V = 1.3 x discharge rate in gpm / rpm V = 0.45 x discharge rate in gpm / rpm
For pump speeds above 100 rpm multiply the above volumes by (Pump rpm / 100). b) Calculation of line shock pressure due to valve closure The magnitude of line shock can be calculated by the following formula:
P= Where,P Vp Sg V v G
Sg × V P × (V − v) 144 × g = Increase in pressure due to shock (psi) = Velocity of pressure wave propagation in pipe (approx 4000 ft/s in small pipes) = Specific gravity of liquid (lb/ft³) = Velocity of liquid in pipe before valve closure (ft/s) = Velocity of liquid in pipe at an interval equivalent to the time that a pressure wave will travel up the pipe and back after the valve starts closing (ft/s) = Acceleration of gravity, 32 (ft/s²)
Example:A 3" line 1000 ft long is carrying water at a pressure of 250 psi and at a velocity of 10 ft/s is suddenly shut-off by a valve closing in 0.3 sec. Assuming that the velocity of the pressure wave in the pipe is approx 4000 ft/s, it would take 0.5 sec for the wave to travel up the pipe and back. This is slower than the time required for the valve to close and consequently v (the velocity thro the valve after the pressure wave has made one complete cycle) is equal to zero.
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN PROCESS DESIGN GUIDELINES
Rev.: A
Date: 08/02/08
Page 24 of 87
P = 62.5 * 4000 * (10 - 0) / (144 * 32.2) = 540 psi Total surge pressure = Original pressure in pipe + shock pressure = 250 + 540 = 790 psi In general water hammer will occur when the total surge pressure exceeds twice the static pressure. Therefore, in the above example, water hammer will occur. c) Calculation of damper volume required to reduce shock Dampers to absorb the shock of fast closing valves are mounted just upstream of the valve. To calculate the minimum surge volume, the following equation may be used.
⎡ R × P2 × ((0.005 × L) − T )⎤ A = 0.004⎢ ⎥ (P2 − P1 ) ⎣ ⎦ Where,
A = Surge volume required (US gpm) R = Pipe flow rate (US gpm) T = Normal closing time of valve in sec. (T=0 for instantaneous closure valves) L = Length of pipe (ft) P1 = Actual flow pressure at the inlet of valve (source pres. - line DP) (psi) P2 = Upper pressure limit which the surge should be limited to in absorbing the decelerating flow on valve closure. This should be set at 1.5 times the static pressure in the line when the valve is closed and the liquid is at rest (psi)
Example: In a 3"-40 sch line 1000 ft long water @ 230 gpm at 10 ft/s is flowing at 250 psi. Fric. Drop is 4.5 psi/100ft. Valve closure time is 0.3 sec. P1 = 250 - (4.5*1000/100) = 205 psi P2 = 250 * 1.5 = 375 psi A = 0.004 * 230 * 375 * (0.005 * 1000 - 0.3) / (375-205) = 9.5 gallons.
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN PROCESS DESIGN GUIDELINES
Rev.: A
Date: 08/02/08
Page 25 of 87
11.
VESSEL SELECTION AND SIZING
11.1
STORAGE TANKS CAPACITY 1) For fixed and roofless tanks the working capacity should be the volume between the top of the suction nozzle and the maximum safe working level in the tank. 2) For floating roof tanks the working capacity should be the volume between the maximum highest safe position for the roof and the minimum allowed position for the roof. 3) For study purposes the working capacity will be multiplied by 1.05 for fixed roof tanks and 1.10 for floating roof tanks to obtain nominal tank capacities for cost estimating purposes.
11.2
TYPE OF ROOF U.S. environmental protection agency requires the following for storage of hydrocarbon in petroleum refineries: 1) Hydrocarbons with a natural vapor pressure of 1.5 psia (0.103 bara) or less at storage temperature may be stored in a freely vented fixed roof tank 2) Hydrocarbons with a natural vapor pressure between 1.5 psia to 11.0 psia (0.103 bara to 0.76 bara) at storage temperature may be stored in floating roof tanks 3) Hydrocarbons with a natural vapor pressure in excess of 11.0 psia (0.76 bara) at storage temperature should be stored in a fixed roof tank with a vapor recovery or refrigeration system. 4) The flammability of the stored material, regulations governing the emission of the vapor to the atmosphere together with cost of the product lost due to evaporation, should be considered.
11.3
TYPE OF STORAGE VESSELS 1) For less than 3.8 m³ (1000 gallons) vertical tanks on legs can be used. 2) Between 3.8 m³ and 38 m³ (1000 to 10,000 gallons) horizontal tanks on concrete supports can be used. 3) Beyond 38 m³ (10,000 gallons) vertical tanks on concrete pads can be used. 4) Liquids with low vapor pressures, tanks with floating roofs can be used. 5) Raw material feed tanks are often specified for 30 days feed supplies 6) Storage tank capacity should be at 1.5 times the capacity of mobile supply vessels. 7) For example, 28.4 m³ (7500 gallon) tanker truck, 130 m³ (34,500 gallon) rail cars 8) Liquid drums are usually horizontal. Gas/Liquid separators are usually vertical 9) Optimum Length/Diameter ratio is usually 3, range is 2.5 to 5
11.4
HOLDUP TIME 1) 2) 3) 4)
Holdup time is 5 minutes for half full reflux drums and gas/liquid separators Design for a 5-10 minute holdup for drums feeding another column For drums feeding a furnace, a holdup of 30 minutes is a good estimate Knockout drum in front of compressors should be designed for a holdup of 10 times the liquid volume passing per minute.
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN PROCESS DESIGN GUIDELINES
Rev.: A
Date: 08/02/08
Page 26 of 87
11.5
VELOCITY CRITERIA 1) Liquid/Liquid separators should be designed for settling velocities of 2-3 inch/min (0.00085 m/s to 0.0013 m/s) 2) Gas velocities in gas/liquid separators
Velocity = k ⋅
ρL
(ρ V
− 1)
Where k =
0.35 with horizontal mesh deentrainers 0.167 with vertical mesh deentrainers 0.1 without mesh deentrainers Velocity is in ft/s
A 6” (150 mm) mesh pad thickness is very popular for such vessels 3) For positive pressure separations, disengagement spaces of 6”-18” (150 mm to 450 mm) before the mesh pad and 12 inches (300 mm) after the pad are generally suitable 11.6
DESIGN CONDITIONS 1) The design pressure for atmospheric storage tanks is +100 mmWC (or full of water) / -50 mmWC 2) Design pressure is 10% or 0.69 bar to 1.7 bar above the maximum operating pressure, whichever is greater, as applicable 3) The maximum operating pressure is taken as 1.7 bar (25 psi) above the normal operation pressure, as applicable 4) For vacuum operations, design pressures are 1 barg (15 psig) to full vacuum 5) For systems with maximum operating temperature (MOT) between -30°C and 345°C, design temperatures is typically MOT + 15°C (to 25°C as applicable). Above this range the margin increases 6) Minimum design temperature is 0°C to 5°C less than the minimum operating temperature 7) Minimum wall thicknesses for maintaining tank structure are: 6.4 mm (0.25 in) for 1.07 m (42 in) diameter and under 8.1 mm (0.32 in) for 1.07 m - 1.52 m (42-60 in) diameter 9.7 mm (0.38 in) for diameters over 1.52 m (60 in) 8) Allowable working stresses are taken as 1/4 of the ultimate strength of the material
11.7
MECHANICAL DESIGN 1) Thickness calculation based on pressure and radius is given by: (Pressure) * (outer Radius) Thickness = (Allowable stress) * (Weld Efficiency) – 0.6 (Pressure) Where : pressure = psig, radius = inch, stress = psi, corrosion allowance = inch ** Weld efficiency can usually be taken as 0.85 for initial design work 2) Guidelines for corrosion allowances are as follows: 0.35 in (9 mm) for known corrosive fluids, 0.15 in (4 mm) for non-corrosive fluids, and 0.06 in (1.5 mm) for steam drums and air
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN PROCESS DESIGN GUIDELINES
Rev.: A
Date: 08/02/08
Page 27 of 87
receivers 3) Generally 1.5 mm to 3.0 mm is considered as the minimum corrosion allowance (CA) for Carbon Steel equipment/piping while Nil corrosion allowance is considered for Stainless Steel and non-metallic equipment/piping 11.8
SIZING EXPLOSION HATCHES
W=
3600 × W ' e
Where W = Required relieving capacity (lb/h) W' = Weight of air and gas in the vessel (lb) e = Time to attain maximum pressure (sec) Time to attain maximum pressure for mixture of gases and air at 1 atm & 150°F is 0.01 sec for H2, 0.045 sec for ethane, 0.056 for propane, 0.06 for hexane and naphtha, 0.0117 for acetylene, 0.06 for benzene, 0.10 for toluene. The required open area, A, to discharge the relieving capacity, W, using a discharge coefficient of 0.8 is:
A =W ×
Z ⋅T / M 245 × P1
Where A = Discharge area (in²) P1 = Relieving pressure (psia) T = Initial temperature (R) Z = Compressibility factor, at P1 & T M = Mixture MW
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN PROCESS DESIGN GUIDELINES
Rev.: A
Date: 08/02/08
Page 28 of 87
12.
HEAT EXCHANGERS Some of the important design tips for heat exchanger design are: 1) Fluids that are corrosive, fouling, scaling, high pressure drop or under high pressure are usually placed in tube side 2) Hot, viscous and condensing fluids are typically placed on the shell side 3) Pressure drops are about 1.5 psi (0.1 bar) for boiling/vaporization and 3-10 psi (0.2-0.7 bar) for other services 4) The minimum approach temperature for shell and tube exchangers is about 20°F (10 °C) for fluids and 10°F (5°C) for refrigerants. 5) Cooling tower water is typically available at a maximum temperature of 90°F (30°C) and should be returned to the tower no higher than 115°F (45°C) 6) Double pipe heat exchangers may be a good choice for areas from 100 to 200 ft2 (9.3-18.6 m2) 7) Spiral heat exchangers are often used to slurry interchangers and other services containing solids 8) Plate heat exchanger with gaskets can be used up to 320°F (160°C) and are often used for interchanging duties due to their high efficiencies and ability to "cross" temperatures. 9) For the heat exchanger equation, Q = UAF (LMTD), use F = 0.9 when charts for the LMTD correction factor are not available 10) Shell and Tube heat transfer coefficient for estimation purposes can be found in many reference books 11) Most commonly used tubes are ¾” (19 mm) outer diameter on a 1” triangular spacing at 16 ft (4.9 m) long 12) A 1 ft (300 mm) shell will contain about 100 ft2 (9.3 m2) A 2 ft (600 mm) shell will contain about 400 ft2 (37.2 m2) A 3 ft (900 mm) shell will contain about 1100 ft2 (102 m2) 13) Typical velocities in the tubes should be 3-10 ft/s (1-3 m/s) for liquids and 30-100 ft/s (9-30 m/s) for gases
12.1
DESIGN MARGINE The following design margins shall be adopted as applicable: 1) Heat exchangers shall be designed with 10% margin on area while using the normal heat duty. 2) While designing heat exchangers, 10% margin on normal duty shall be considered and the exchanger shall be designed with minimum over design factor (say 2% to 5%) on area. 3) If the design duty (10% to 15% over design factor) is used for sizing the heat exchangers, then the exchanger shall be designed with minimum over design factor (say 2% to 5%) on area.
12.2
EVAPORATORS 1) When the boiling point rise is appreciable, the economic number of effects in series with forward feed is 4-6. 2) When the boiling point rise is small, minimum cost is obtained with 8-10 effect in series.
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN
Rev.: A
PROCESS DESIGN GUIDELINES
Date: 08/02/08
Page 29 of 87
12.3
HEAT EXCHANGER DETAILS 1) TEMA shell types:
TEMA X
TEMA G
TEMA H
TEMA J21
TEMA J12
TEMA E
TEMA K ACU
2) Baffle cut orientation; parallel, perpendicular and 45° with reference to shell nozzle Parallel 0°
45°
Perpendicular 90°
3) Tube length : Length up to the tangent point of the outermost tube for U-tubes. The length in the tube sheets should be included. Effective area is excluding the tube sheet. U-tube exchanger
Overall tube length Total tube length
4) Shell inlet nozzle location Horizontal
Vertical code 0
code 1
code 2
5) U-tube nozzle location
Code 0, normal
Code 1, in front of
Code 2, behind U
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN PROCESS DESIGN GUIDELINES
Rev.: A
Date: 08/02/08
Page 30 of 87
6) Shell side inlet nozzle location: Normally it is assumed to be at the tube channel end.
Code N, shell inlet at tube end
Code Y, shell inlet at tube channel end
Distance from tangent point to last baffle: Generally the last baffle is placed at the at the tangent point of the U-tube 7) Tube arrangement Triangular Pitch
Rotated Square Pitch
Rotated Triangular Pitch
Square Pitch
Flow 30°
45°
Pitch
Pitch
60°
Pitch
90°
Pitch
8) Impingement device: An impingement plate will be added if the inlet nozzle pV² value is greater than TEMA standards (1500 lb/ft s² or 2232 kg/m s²) 9) Adding a shell in parallel: If pressure drop limitations are not met at the maximum shell diameter permitted, a unit can be added in parallel. 10) Adding a shell in series: If a pure counter flow exchanger has been selected, a shell can be added in series only if the required heat duty can not be achieved at the maximum permitted shell diameter. For multi-tube pass exchangers shells in series can be added if the F-factor is less than 0.7 11) Temperature cross: When the outlet temp of the cold fluid is higher than the outlet temp of hot fluid, it is called the temperature cross. For small exchangers, temperature cross does not have much effect on the type of shells. But for large exchangers shells in series to be used. 12) Baffle design: Baffle window cut should be between 17% to 35%. Baffle spacing should be 20% to 100% of shell ID. For no-tube-in-window exchangers, the ratio of window velocity to cross-flow velocity should normally be 2 to 3. For double segmental baffles (for low pres drop service), baffle spacing should not be too small to avoid ineffective shell side flow patterns. 13) No-tube-in-windows baffle cut design: The baffle cut shall be limited between 15 - 30% of shell diameter 14) Baffle cut out of window: The portion of the baffle which is continuous into the window acts as a sealing strip in the window to force the fluid into the bundle. Continuous baffles should be considered for pull-through floating heads.
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN
Rev.: A
PROCESS DESIGN GUIDELINES
Date: 08/02/08
Page 31 of 87
12.4
HTRI SHELL SIDE FLOW FRACTIONS This is for shells with single phase fluid flow. The shell flow is broken down into 5 major streams. E
Window
A
A
B
C
C
F
C
A
A
Baffle
E
1) B-stream : Main cross flow stream through the bundle. B should be at least 60% of the total flow for turbulent flow and 40% for laminar flow. If the baffle spacing is too narrow, more flow will be forced into the A, C and E streams, thereby decreasing the heat transfer. 2) C-stream : Bundle to shell cross flow bypass stream. C should not normally exceed 10%. Additional sealing strips can be incorporated to decrease this flow fraction. Although this stream is partially effective for heat transfer, a high C-stream flow fraction, especially for pure cross flow shells ("X"), can lead to a severe delta correction to the mean temp difference 3) F-stream : Tube pass partition bypass stream. F should not normally exceed 10%. Additional seal rods can be incorporated to decrease this flow fraction. Although this stream is partially effective for heat transfer, a high C-stream flow fraction, is not recommended. To block the F-stream flow fraction, program assumes one seal rod of a diameter equal to the tube diameter for each 6 tube rows of cross flow in the exchanger 4) A-stream : Tube-to-baffle hole leakage stream. A-stream is large in narrow baffle spacing where large TEMA clearances apply. However, the A-stream is fairly effective thermally. It will decrease for multi-segmental baffles. Fouling layers might seal this A-stream. The design should be examined by giving a zero tube-to-baffle clearance and the built-up fouling layer thickness for a safe design from a pres drop stand point. 5) E-stream : Baffle-to-shell leakage stream. E-stream is highly ineffective thermally because it does not contact the heat transfer surface; but, since it mixes poorly with the other streams, it can cause distortions of the temp profile. If E-stream is more than 15%, double segmental baffle or other modifications should be tested. If E-stream causes (ST program only) low delta correction factor (< 0.8), corrective action is required. •
F-stream seal rods: Allocate one seal rod of the tube diameter for each 6 rows in cross flow in the exchanger
•
Sealing strips: These are metal strip or rod placed between the shell and the bundle which has the effect of forcing the bundle bypass C-stream back into the bundle.
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN PROCESS DESIGN GUIDELINES
Rev.: A
Date: 08/02/08
Page 32 of 87
12.5
SHELL SIDE ANNULAR DISTRIBUTOR
Clearance Slot area
Length
Thermal correction factor "F" = (TUBE) x (BAFFLES) x (F/G) x (HOT/COLD); where TUBE is the uncorrected F-factor based on the no of tube passes, shell style and temp BAFFLES is the correction when there are few baffles. (F/G) is the correction for thermal leakage through longitudinal baffle for TEMA "F","G","H" shells. (HOT/COLD) is the correction for nonconstant overall h.t.coeff due to diff in the h.t.coeff at the hot and cold ends. Effective MTD = (LMTD) x F x (DELTA): where DELTA is the profile distortion due to the Eand C-stream leakage 12.6
SHELL SIDE HEAT TRANSFER LIMIT If there is spare shell side pres drop available, the shell side coefficient can be increased by various methods: 1) Changing the shell type to "F" or "G" can increase the shell side velocity and h.t.coefficient. But the mean temperature difference may increase. 2) Reducing the tube pitch 3) Decreasing the tube size to accommodate more tubes in a smaller shell 4) Considering finned tubes 5) If DELTA is lower than 0.85, adding a shell in series gives best result, or using sealing strips might improve the performance
12.7
TUBE SIDE HEAT TRANSFER LIMIT If there is spare tube side pressure drop available, the tube side coefficient can be increased by various methods: 1) Changing the tube length 2) Decreasing tube diameter 3) When in laminar flow switching the tube side fluid to shell side usually results in a more efficient design 4) Increasing the tube pitch, gives less tubes in the given shell ID
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN
Rev.: A
PROCESS DESIGN GUIDELINES
Date: 08/02/08
Page 33 of 87
12.8
LIQUID DRIVING HEAD FOR THERMOSIPHON EXCHANGERS
C
Main pipe
H Liquid Driving Head (Static Head)
G
D
B
C F
E Nozzle pipe
A
Header pipe Main pipe
1) For horizontal thermosiphon reboilers, most recirculating type feed systems can be designed with kettle type since the height of the outlet piping entering the column is above the liquid level in the column as shown in figure. 2) For thermosiphon reboiler systems for which the reboiler outlet piping enters the distillation column at a height below the liquid level in the trap-out tray as shown in figure the piping should be checked to ensure that the liquid level does not cover the exit nozzle. 3) For horizontal thermosiphon reboiler designs, the reboiler exit weight fraction should be limited to 0.5 to avoid tube wall dry out. 4) For an effective design, most of the available loop pressure drop is used across the reboiler. As a rule of thumb, this should be around 60 - 70%. However, the inlet and outlet piping design may change this requirement.
Recirculating Feed System
Once through Feed System
Vapor+Liquid
Vapor+ Liquid
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN PROCESS DESIGN GUIDELINES
Rev.: A
Date: 08/02/08
Page 34 of 87
13.
COLUMNS & TOWERS
13.1
GENERAL DESIGN RULES 1) For ideal mixtures, relative volatility can be taken as the ratio of pure component vapor pressures 2) Tower operating pressure is most often determined by the cooling medium in condenser or the maximum allowable reboiler temperature to avoid degradation of the process fluid 3) For sequencing columns: • Perform the easiest separation first (least trays and lowest reflux) • If relative volatility or feed composition varies widely, take products off one at time as the overhead • If the relative volatility of components does vary significantly, remove products in order of decreasing volatility • If the concentrations of the feed vary significantly but the relative volatility does not, remove products in order of decreasing concentration. 4) The three most common types of trays are valve, sieve, and bubble cap. Bubble cap trays are typically used when low-turn down is expected or a lower pressure drop than the valve or sieve trays can provide is necessary 5) Bubble cap trays are used only when a liquid level must be maintained at low turn down ratio; they can be designed for lower pressure drop than either sieve or valve trays. 6) The optimum Kremser absorption factor is usually in the range of 1.25 to 2.00
13.2
TRAY COLUMNS/TOWERS 1) The most economic number of trays is usually about twice the minimum number of trays 2) The minimum number of trays is determined with the Fenske-Underwood Equation 3) Typically, 10% more trays than calculated are specified for a tower 4) Tray spacing’s should be from 18” to 24” (450 to 600 mm), with accessibility in mind 5) Peak tray efficiencies usually occur at linear vapor velocities of 2 ft/s (0.6 m/s) at moderate pressures, or 6 ft/s (1.8 m/s) under vacuum conditions. 6) Pressure drop per tray is of the order of 3” of water or 0.1 psi (0.007 bar) 7) Tray efficiencies for distillation of light hydrocarbons and aqueous solutions are usually in the range of 60-90% while gas absorption and stripping typically have efficiencies closer to 10-20% 8) Sieve tray holes are 0.25 to 0.50 in. diameter with the total hole area being about 10% of the total active tray area 9) Valve trays typically have 1.5” diameter holes each with a lifting cap. 12-14 caps/square foot of tray is a good benchmark. Valve trays usually cost less than sieve trays 10) The most common weir heights are 2” and 3” (50 to 80 mm) and the weir length is typically 75% of the tray diameter 11) For towers that are at least 3 ft (0.9 m) in diameter, 4 ft (1.2 m) should be added to the top for vapor release and 6 ft (1.8 m) should be added to the bottom to account for the liquid level and reboiler return 12) Limit tower heights to 175 ft (53 m) due to wind load and foundation considerations 13) The Length/Diameter ratio of a tower should be no more than 30 and preferably below 20 14) Liquid redistributors are needed every 5-10 tower diameters with pall rings but at least every 20 feet (6.1 m). The number of liquid streams should be 3-5 /ft² in towers larger than 3 ft dia, and more numerous in smaller towers 15) A rough estimate of reboiler duty as a function of tower diameter is given by: Q = 0.5 D² for pressure distillation Q = 0.3 D² for atmospheric distillation Q = 0.15 D² for vacuum distillation
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN PROCESS DESIGN GUIDELINES
Rev.: A
Date: 08/02/08
Page 35 of 87
Where Q = Reboiler duty (Million Btu/h) D = Tower diameter (ft) 13.3
REFLUX DRUMS 1) The most economic reflux ratio usually is between 1.2 Rmin and 1.5 Rmin 2) Reflux pumps should be at least 25% over designed 3) Reflux drums are almost always horizontally mounted and designed for a 5 min holdup at half of the drum's capacity. A take off pot for second liquid phase such as water in hydrocarbon systems, is sized for a linear velocity of that phase of 0.5 ft/s (0.15 m/s) with minimum diameter of 16” (400 mm)
13.4
PACKED TOWERS 1) Packed towers almost always have lower pressure drop than comparable tray towers 2) Packing is often retrofitted into existing tray towers to increase capacity or separation 3) For gas flow rates of 500 ft³/min (14.2 m³/min) use 1 in (2.5 cm) packing, for gas flows of 2000 ft³/min (56.6 m³/min) or more, use 2 in (5 cm) packing 4) Ratio of tower diameter to packing diameter should usually be at least 15 5) Due to the possibility of deformation, plastic packing should be limited to an unsupported depth of 10-15 ft (3-4 m) while metallic packing can withstand 20-25 ft (6-7.6 m) 6) Liquid distributor should be placed every 5-10 tower diameters (along the length) for pall rings and every 20 ft (6.5 m) for other types of random packing 7) For redistribution, there should be 8-12 streams per sq. foot of tower area for tower larger than 3 feet in diameter. They should be even more numerous in smaller towers 8) Packed columns should operate near 70% flooding 9) Height Equivalent to Theoretical Stage (HETS) for vapor-liquid contacting is 1.3-1.8 ft (0.40.56 m) for 1 in pall rings and 2.5-3.0 ft (0.76-0.90 m) for 2 in pall rings
13.5
DESIGN PRESSURE DROPS Service Absorbers and Regenerators Non-Foaming Systems Moderate Foaming Systems Fume Scrubbers Water Absorbent Chemical Absorbent Atmospheric or Pressure Distillation Vacuum Distillation Maximum for Any System
Pressure drop (in H2O/ft packing) 0.25 - 0.40 0.15 - 0.25 0.40 - 0.60 0.25 - 0.40 0.40 - 0.80 0.15 - 0.40 1.0
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN PROCESS DESIGN GUIDELINES
Rev.: A
Date: 08/02/08
Page 36 of 87
14.
COMPRESSORS AND VACUUM EQUIPMENT The following chart is used to select the type of compressor:
1) Fans should be used to raise pressure about 3% (12” or 300 mm water), blowers to raise to less than 2.75 barg (40 psig), and compressors to higher pressures 2) The theoretical reversible adiabatic power is estimated by:
Power = m ⋅ Z ⋅ R ⋅ T1 ×
[( P2 / P1 ) × a] − 1 a
where: T1 = Inlet temperature; R = gas constant; Z1 = Compressibility; m = molar flow rate a = (k − 1) / k
k = Cp / Cv T2 = T1 × ( P2 / P1 ) × a T2 = Outlet temperature for adiabatic reversible flow 3) Exit temperatures should not exceed 204°C (400°F) 4) For diatomic gases (Cp/Cv = 1.4) this corresponds to a compression ratio of about 4 5) Compression ratios should be about the same in each stage for a multistage unit,
⎛P Ratio = ⎜⎜ n ⎝ P1
⎞ ⎟⎟ ⎠
1/ n
with n stages.
6) Efficiencies for reciprocating compressors are: - 65% at compression ratios of 1.5 - 75% at compression ratios of 2.0 - 80-85% at compression ratios between 3 and 6
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN PROCESS DESIGN GUIDELINES
Rev.: A
Date: 08/02/08
Page 37 of 87
7) Efficiencies of large centrifugal compressors handling 2.8 to 47 m³/s (6000-100,000 acfm) at suction is about 76-78% 8) Flash gas compressors typically have an overall compressor ratio in the range 5 to 10 9) For compressors, the brake horsepower per stage can be determined from k −1 ⎤ ⎡ ⎞ ⎛ k ⎞ ⎢⎛ Pd ⎞ k ⎥ ⎟⋅⎜ BHP = 0.0857 ⋅ (Z av ) ⋅ (Z s ) ⎟ ⎝ k − 1 ⎟⎠ × ⎢⎜⎜ P ⎟⎟ − 1⎥ ⎠ ⎥⎦ ⎢⎣⎝ s ⎠ Where BHP = Brake horsepower per stage Qg = Volume of gas (MMscfd) Ts = Suction temperature (R) Zs = Suction compressibility factor Zd = Discharge compressibility factor Zav = (Zs + Zd) / 2 E = Efficiency % High speed reciprocating units – use 70% Low speed reciprocating units – use 78% Centrifugal units - use 75% K = Ratio of specific heat, Cp/Cv Ps = Suction pressure of stage (psia) Pd = Discharge pressure of stage (psia) 1 k
14.1
k −1 k
⎛ Q g ⋅ Ts ⋅ ⎜⎜ ⎝ E
VACUUM PUMPS 1) Reciprocating piston vacuum pumps are generally capable of vacuum to 1 torr (1 mmHg absolute) 2) Rotary piston types can achieve vacuums of 0.001 torr 3) Single stage Jet ejectors are capable of vacuums up to 100 torr, 2-stages to 10 torr, 3stages to 1 torr, and 5-stages to 0.05 torr 4) A three stage ejector requires about 100 lb steam/lb air (100 kg steam/kg air) to maintain a pressure of 1 torr 5) Air leakage into vacuum equipment can be approximated as follows:
Leakage = k × V 2 / 3 Where k = 0.20 for P >90 torr, 0.08 for 3 < P < 20 torr, and 0.025 for P < 1 torr V = equipment volume (ft³) Leakage = air leakage into equipment (lb/h)
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN PROCESS DESIGN GUIDELINES
Rev.: A
Date: 08/02/08
Page 38 of 87
15.
SAFETY SYSTEM & PSV DESIGN
15.1
SAFETY SYSTEM CONFIGURATION 1) For direct discharge of fluids to atmosphere, they should be in the vapor state, below their auto ignition temperature, and should meet one of the following requirements:• Flammable vapors of MW less than 28.9 (MW of air). • Flammable vapors heavier than air with MW less than 70 but with minimum discharge velocity of 500 fps (152 m/s), based on maximum capacity of the relief valve • Vapors of any MW that are non-flammable, non-toxic and non-condensable 2) In all other cases relieved fluids should be disposed to a flare system. 3) Boilers having more than 47 m² of water heating surface or electric boilers having a power input of more than 500 kW shall have 2 or more PSV. 4) For electric boilers the minimum relieving capacity shall be 1.6 kg/hr/kW input 5) For Boilers, if additional PSV are used the highest pressure setting shall not exceed the MAWP by more than 3%. When multiple PSV are installed, the difference between the highest and lowest set pressure should not be greater than 10% of the highest set pressure. 6) For Pressure vessels, if additional PSV are used the highest pressure setting shall not exceed 105% of MAWP. For multiple PSV, one of which is installed for fire exposure only, this particular valve may be set at a pressure not exceeding 110% of MAWP 7) Fire exposure protection of vessels by water spray @ 0.05 - 0.2 gpm/ft² of total vessel area. Fire exposure is considered for a wetted area up to 25 ft (7.62 m) from grade. Average or normal liquid level to be considered for vessels, High liquid level for columns, surge drums and KO drums. 50% of storage tank height or 25 ft (7.62 m) from grade, which ever is higher. 8) For flare fire relief load, consider plot area between 2000-5000 ft² (186-465 m²). Use 2500 ft² (232 m²) for a paved drained surface in a plant where NFPA required fire fighting equipment is available. 9) If fire proofing insulation is not provided, environmental factor (F) should be = 1.0
15.2
SAFETY VALVE BACK PRESSURE 1) Super imposed back pressure = pressure at PSV discharge before valve opening 2) Builtup back pressure = pressure at PSV discharge header after valve opening. 3) Conventional PSV shall be used when the superimposed back pressure varies over a range not exceeding 10% of the set pressure (gauge). However the performance of PSV at builtup back pressure to be studied by the vendor before selecting the PSV 4) Balanced type PSV can be used for >10% superimposed back pressure. However when builtup back pressure exceeds 30-50% of set pressure, capacity of PSV for vapors and gases starts to fall below the theoretical capacity. With liquids, the capacity reduction starts at 15% of set pressure. The fall in capacity depends on overpressure, type and make of PSV. 5) Pilot operated relief valves are used when operating pressure is very close to set pressure or when set pressure is below 10-15 psig (0.7 – 1.0 barg)
15.3
SAFETY VALVE RELIEF LOAD CALCULATION PSV’s relief load can be calculated by considering the following major failure scenarios:
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN PROCESS DESIGN GUIDELINES
Rev.: A
Date: 08/02/08
Page 39 of 87
1) 2) 3) 4) 5) 6) 7) 8)
Blocked outlet Gas blow by case Inadvertent valve opening Reflux failure Tube rupture Fire case Reverse flow Thermal relief case.
15.3.1 Blocked Outlet Blocked outlet case can occur when the control valves at the outlet of the vessel close at the same time due to the stoppage of instrument air supply or plant shut down. The relief quantity shall be the design flow rate to the vessel. 15.3.2 Fire Case The amount of heat absorbed by a vessel open to fire is markedly affected by the type of fuel feeding the fire. The following equivalent formulas are used to evaluate the condition where there are prompt firefighting efforts and drainage of flammable materials away from the vessels are available: q = 21000 × F × A −0.18
Q = 21000 × F × A 0.82 Where adequate drainage & firefighting equipment do not exist, the following equation should be used: Q = 34500 × F × A 0.82 Where q = Average unit heat absorption (Btu/h/ft² of wetted surface) Q = Total heat absorption (input) to the wetted surface (Btu/h) F = Environmental factor (Values for various type of installation are given in Table 5 in API 521 page 25) A = Total wetted surface (ft²) The discharge area for pressure relief devices on vessels containing super critical fluids, gases or vapors exposed to open fires can be estimated by using the following equation.
A= F'=
F '× A' P1
0.1406 ⎛⎜ (Tw − T1 )1.25 C ⋅ K D ⎜⎝ T1 0.6506
⎞ ⎟ ⎟ ⎠
The relief load can calculated directly, in pounds per hour.
⎛ (T − T1 )1.25 ⎞ ⎟ W = 0.1406 × M ⋅ P1 × ⎜ A' w 1.1506 ⎜ ⎟ T 1 ⎝ ⎠
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN PROCESS DESIGN GUIDELINES
Rev.: A
Date: 08/02/08
Page 40 of 87
Where A = Effective discharge area of the valve (in²) A’ = Exposed surface area of the vessel (ft²) F’ = PSV Factor (= 0.01 min. or 0.045 if value is not known) P1 = Upstream relieving pressure (psia) P1 = PSV set pressure + allowable over pressure + atmospheric pressure C = Cp/Cv KD = Coefficient of discharge (=0.975 max) M = Molecular weight of the gas TW = Vessel wall temperature (R) (for CS vessels = 1100°F) T1 = Gas temperature (R) at the upstream relieving pressure, determined from the following relation.
T1 = ( P1 / Pn ) × Tn Where Pn = Normal operating gas pressure (psia) Tn = Normal operating gas temperature (R) Air Coolers For air coolers heat absorption equation becomes Q = 21000 × F × A1.0 Total bare tube area to be considered instead of finned area, since fins are burned out in the first few minutes of fire. Water spray nozzles are some times mounted below the tubes in case of fire. API 520 recommends a minimum of 0.05 - 0.2 gpm/ft² water spray rate. 15.3.3 Inadvertent Valve Opening
CV =
W N 6 × FP × Y × X × P1 × γ 1
Y =1−
X 3 × FK × X K
Where W = Mass to be relieved (kg/h)
N 6 = Constant (27) FP = Geometric factor Y = Downstream pressure (bara) X = Ratio of pressure drop P1 = Upstream pressure (bara)
γ 1 = Upstream gas density (kg/m³)
15.3.4 Thermal Expansion As per API RP 521, the relieving capacity requirements for hydraulic expansion cases will be very small as relieving fluid will be liquid. Therefore ¾” X 1” nominal pipe size relief valve is commonly used
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN PROCESS DESIGN GUIDELINES
Rev.: A
Date: 08/02/08
Page 41 of 87
Cold fluid blocked in and continuous heat input from hot fluid. Two conditions can occur: Cold fluid remains liquid and expands WL = B × Q / c Cold fluid vaporizes
WV =
(T1 − Tbp ) (T1 − Tav )
×Q/ L
Where WL, Wv = weight of liquid or vapor relieved (lb/h) Q = Normal exchanger duty (Btu/h) L = Latent heat of vaporization at relieving conditions (Btu/lb) T1 = Hot side inlet temperature (°F) Tbp = Cold side boiling temperature at relieving pressure (°F) Tav = Average of inlet and outlet temperature of cold side during normal operation (°F) c = Specific heat of cold medium (Btu/lb °F ) B = Coefficient of expansion of cold medium (1/°F) Fluid Oil Oil Oil Oil
°API 3 to 34.9 35 to 50.9 51 to 63.9 64 to 78.9
Expansion Coef. 0.0004 0.0005 0.0006 0.0007
Fluid Oil Oil Oil Water
°API 79 to 88.9 89 to 93.9 =>94
Expansion Coef. 0.0008 0.00085 0.0009 0.0001
15.3.5 Tube failure and leakage Tube failure is considered a viable contingency when the design pressure of the low pressure side is less than 2/3 rd (=67%) of design pressure of high pressure side. However, if the high pressure side of the exchanger operates at 1000 psig (69 barg) or more and contains a vapor or liquid that can flash or result in vaporization of liquid on the low pressure side, complete tube failure should be considered, regardless of the pressure differential. Install a PSV, if the piping and downstream equipment on the low pressure side do not have the capacity to handle material leaked from high pressure side without exceeding 110% of the equipment design pressure. A tube failure is considered to be a sharp break in one tube. The high pressure fluid flows through both openings, which is equal to twice the cross section area of a single tube. Following equation can be used to calculate the flow from high to low pressure sides.
WL = 2 × 1.256 × d 2 × C × (∆P ×ρ Where WL = Relieving rate. (kg/h)
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN PROCESS DESIGN GUIDELINES
Rev.: A
Date: 08/02/08
Page 42 of 87
d = Tube inside diameter (mm) C = Orifice coefficient ∆P = Differential pressure (Operating pressure of high pressure side (bara) Relieving pressure of PSV (1.1 times set pressure of PSV) (bara) ρ = Density of high pressure fluid (kg/m³) For a discharge coef. of 0.7, Cp/Cv=1.33 and twice the cross sectional area of one tube
q L = 41.8 × d 2 ×
P1 − P2 Sg
qV = 385 × d 2 × P1 ×
M Z ×T
Where qL, qv = Quantity of liquid (gpm) or vapour (lb/h) d = tube inside diameter (inch) M = MW P1 = normal high pressure side (psig) Sg = specific gravity P2 = 1.1 times low pressure design pressure (psig) Z = compressibility factor T = vapor temperature (°R) at operating conditions If the calculated discharge exceeds the normal total flow in the high pressure side, the latter flow should be used. Possible flash of liquid to vapor shall be taken into account due to both pressure reduction and mixing of a volatile fluid with a hot fluid. Valves on low pressure side provided only for isolation may be assumed fully open, control valves in a position equivalent to the minimum normal flow unless the valve could automatically close due to the emergency situation. 15.4
SETTLE-OUT PRESSURE In high pressure oil & gas handling facilities the compressor suction KO drum has to be designed based on settle-out pressure. In case of compressor trip the suction and discharge shutdown valves will close and the gas from high pressure discharge side will flow through the recycle control valve to the suction side. A new settle-out condition is reached. The settle-out pressure is calculated by the following formula: ( P × V ) + ( P2 × V2 ) Ps = 1 1 V1 + V2 Where Ps P1 P2 V1 V2
= Settle-out pressure (barg) = Suction side pressure before settle-out (barg) = Discharge side pressure before settle-out (barg) = Suction side volume (including KO drum, piping, etc) (m³) = Discharge side volume (including KO drum, piping, etc) (m³)
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN PROCESS DESIGN GUIDELINES
Rev.: A
Date: 08/02/08
Page 43 of 87
16.
PIPE SIZING The major criteria for line sizing are velocity and pressure drop 1) Liquid lines should be sized for a velocity of (5+D/3) ft/s and a pressure drop of 2.0 psi/100 ft of pipe at pump discharge. At the pump suction, size for (1.3+D/6) ft/s and a pressure drop of 0.4 psi/100 ft of pipe. **D = pipe diameter (inch) 2) Steam or gas lines can be sized for 20D ft/s and pressure drops of 0.5 psi/100 ft of pipe 3) Limits on superheated, dry steam or gas line should be 61 m/s (200 ft/s) and a pressure drop of 0.1 bar/100 m (0.5 psi/100 ft) of pipe. Saturated steam lines should be limited to 37 m/s (120 ft/s) to avoid erosion. 4) For turbulent flow in commercial steel pipes, use the following:
M ⋅ µ 0.2 ⋅ 20000 × D ⋅ ρ Where:∆P = Frictional pressure drop (psi/100 ft) M = Mass flow (lb/h) µ = Viscosity (cP) ρ = Density (lb/ft³) D = ID of pipe (inch) For smooth heat exchanger steel tubes replace 20000 with 23000 ∆P =
5) For two phase flow, an estimate often used is Lockhart and Martinelli. First, the pressure drops are calculated as if each phase exist alone in the pipe, then 0.5
⎛ ∆P ⎞ X ⋅ = ⋅⎜⎜ L ⎟⎟ ⎝ ∆PG ⎠ Now the total pressure drop can be calculated by one of the following:
∆PTOTAL = YL ⋅ ∆PL ⋅ ⋅or ⋅ YG ⋅ ∆PG Where YL = 4.6 ⋅ X −1.78 + 12.5 ⋅ X −0.68 + 0.65 YG = X 2 ⋅ YL 6) Control valves require at least 0.7 bar (10 psi) pressure drop for sufficient control 7) Flange ratings include 10, 20, 40, 103, and 175 bar (150, 300, 600, 1500, and 2500 psig) 8) Globe valves are most commonly used for gases and when tight shutoff is required. Gate valves are common for most other services. 9) Screwed fitting are generally used for line sizes 2 inches and smaller. Larger connections should utilize flanges or welding to eliminate leakage. 10) Pipe Schedule Number = 1000 P/S (approximate) where P is the internal pressure rating in psig and S is the allowable working stress of the material is psi. Schedule 40 is the most common. 11) About 15% margin in flow should be considered in general for line sizing 12) For fuel oil and heating oil, lines to be sized for 25% margin on flow 13) Pump suction lines should be checked for NPSH requirement. More critical if the liquid is at its saturation temperature
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN PROCESS DESIGN GUIDELINES
Rev.: A
Date: 08/02/08
Page 44 of 87
14) 100% margin is desirable for pump suction lines 15) Pump suction line is generally one (or) two sizes higher than the discharge line 16) Liquid overflow line size shall be at least one size higher than the largest liquid inlet line to that equipment 17) Most of the time, the line size is same as that of the vessel nozzle 18) For pumps and compressors, the connected pipeline (inlet and outlet) sizes are generally higher than those of the equipments 19) Safety valve inlet line sizing is based on a pressure drop corresponding to about 3% of the set pressure (in absolute) 20) Minimum velocity of 1 m/sec is preferred for water services and liquids carrying suspended solid 21) The pressure downstream of PSV, shall not exceed 10% of the PSV set pressure for conventional safety valves. For balanced type valves it is preferable to limit this pressure to 40% of the set pressure 22) For free draining line, sizing should be less than 1 m/sec velocity
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN PROCESS DESIGN GUIDELINES
Rev.: A
Date: 08/02/08
Page 45 of 87
Recommended Velocity and Pressure Drop for Process Liquids (not Water) SL. NO.
Type of Service
Recommended Velocity (m/s) 1.5 - 4.5
Max. Allowable Pressure drop (bar/km) 9.0
1.2 - 1.5
---
3.0 - 4.5 1.8 - 3.0 1.5 - 2.5
-------
1.
General recommendation
2.
0.6 - 1.8 1.2 - 2.4
1.1 2.2
6.
Laminar flow Turbulent flow Liquid density (kg/m³) 320 800 1600 Pump suction Boiling liquid Non-boiling liquid Pump discharge (m³/h) < 57 57 - 159 > 159 Liquid flow on gravity
1.8 - 2.4 2.4 - 3.0 3.0 - 4.5 0.9 – 2.4
13.5 9.0 4.5 0.9
7.
Vessel / Column Inlet
1.2 - 1.8
---
3.
4.
5.
8.
Outlet (+vortex breaker)
9.
Outlet to reboilers
10.
Liquid from total condenser
1.0
0.9
0.3 - 1.2
0.33
0.9 - 1.8
1.1
11.
Liquid to/from chillers
1.2 - 1.8
---
12.
Refrigerant liquid
0.6 - 1.2
0.9
13.
Self venting lines (max velocity)
0.18
---
NOTE : The velocity higher than 3.0 m/sec for any line should be used only after due consideration for velocity limitations.
Recommended Velocity and Pressure Drop for Steam Services
Type of Service Steam Headers : Saturated Superheated
Recommended Velocity (m/s)
Max. Allowable Pressure Drop (bar/km)
60 (max) 75 (max)
Steam Pressure 0 - 3.5 barg 3.5 – 10 barg 10 – 20 barg 20 barg and above Process steam piping (saturated)
30 - 50
0.57 1.1 2.2 3.4 ---
Boiler steam piping (superheated)
35 - 100
6.8
Steam lines to Turbine (superheated)
35 - 100
6.8
Steam lines to Turbine exhaust header
---
3.4
Turbine exhaust to header ( > atm)
---
1.1
3.8 - 4.6
---
Driving steam to pumps & recip. equip.
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN PROCESS DESIGN GUIDELINES
Rev.: A
Date: 08/02/08
Page 46 of 87
Recommended Velocity and Pressure Drop for Water (Condensate) Recommended Velocity (m/s) 0.6 - 5.0
Max. Allowable Pressure drop (bar/km) 4.5
2.
Pipe diameter < 1” 2” 4” 6” 8” 10” 12” 16” 20” and above Pump suction
0.6 - 0.9 0.9 - 1.3 1.5 - 2.0 2.0 - 2.7 2.4 - 3.0 3.0 - 3.6 3.0 - 4.3 3.0 - 4.6 3.0 - 4.8 1.2 - 2.1
4.5 4.5 4.5 4.5 4.5 4.5 4.5 4.5 4.5 ---
3.
Pump discharge
1.5 - 3.0
---
4.
Boiler feed water
2.4 - 4.5
---
5.
Condensate from condenser
0.9 - 1.5
---
6.
Process water Cooling water (once thro, brackish/sea) Cooling water (circulating, clean) Water drain lines (gravity flow)
0.6 - 1.5
5.5
3.6 - 4.9
4.5
1.5 - 3.0
5.0
1.2 - 2.1
---
SL. NO.
Type of Service General service
1.
7. 8. 9.
Recommended maximum Velocity for Special Liquids or Materials Sl. No.
Type of Service
Recommended Maximum Velocity (m/s)
Carbon steel pipe with 1.
2. 3. 4. 5.
Phenolic water Concentrated H2SO4 Salt water Caustic solution (low temp.) Stainless steel pipe with CO2 -rich amine liquid Cement pipe with salt water Coal Tar enamel lined pipe with salt water Plastic pipe or rubber lined pipe with Liquid in general Liquid with suspended solids
0.9 1.2 1.8 1.2 3.0 4.5 4.5 3.0 0.9 (min. velocity)
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN PROCESS DESIGN GUIDELINES
Rev.: A
Date: 08/02/08
Page 47 of 87
Recommended Velocity and Pressure Drop for Vapors (not Steam) Type of Service General recommendation Pressure 0.7 barg and less Up to 3.5 barg 3.5 - 10 barg 10 - 14 barg 14 - 34 barg 34 barg and above Vessel / Column outlet Pressure 0.7 barg and less Up to 3.5 barg Above 3.5 barg
Recommended Velocity (m/s)
Max. Allowable Pressure Drop (bar/km)
-------------
0.22 0.34 0.68 1.3 3.4 4.5
38 - 60 18 - 30 12 - 15
0.11 - 0.22 0.45 - 1.1 0.45 - 1.1
Compressor suction
---
1.1
Compressor discharge
---
2.2
Refrigerant suction
5.0 - 11.0
---
Refrigerant discharge
11.0 - 18.0
---
---
1.1
General gas lines Vapor lines (excluding flare lines)
Maximum ρV2 (1) (kg/m.s2)
Continuous operation: P ≤ 20 bar g 20 < P ≤ 50 bar g 50 < P ≤ 80 bar g 80 < P ≤ 120 bar g P > 120 bar g
6,000 7,500 (2) 10,000 15,000 20,000
Discontinuous operations e.g.: Compressor anti-surge P ≤ 50 barg 50 < P ≤ 80 barg P > 80 barg
10,000 (2) 15,000 25,000
(1) Gas density in kg/m³, V = gas velocity in m/s (2) Value could be increased but not exceeding 15,000 in case of debottlenecking
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN PROCESS DESIGN GUIDELINES
Rev.: A
Date: 08/02/08
Page 48 of 87
17.
SELECTION OF VALVES Some of the most common types of valves are:
17.1
GATE VALVES Best suited control: Quick opening Recommended uses: 1. Fully open/closed, non-throttling 2. Infrequent operation 3. Minimal fluid trapping in line Applications: Oil, gas, air, slurries, heavy liquids, steam, non-condensing gases, and corrosive liquids Advantages High Capacity Tight shut-off Low cost Little resistance to flow
17.2
Disadvantages Poor control Cavitate at low pressure drops Cannot be used for throttling
GLOBE VALVES Best suited control: Linear and equal percentage Recommended uses: 1. Throttling service/flow regulation 2. Frequent operation Applications: Liquids, vapors, gases, corrosive substances, slurries Advantages Efficient throttling Accurate flow control Available in multiple ports
17.3
Disadvantages High pressure drop More expensive than other valves
BALL VALVES Best suited control: Quick opening, linear Recommended uses: 1. Fully open/closed, limited-throttling 2. Higher temperature fluids Applications: Most liquids, high temperatures, slurries Advantages High Capacity Tight sealing with low torque Low cost Low leakage and maintenance
Disadvantages Poor control Prone to Cavitation
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN PROCESS DESIGN GUIDELINES
Rev.: A
Date: 08/02/08
Page 49 of 87
17.4
BUTTERFLY VALVES Best suited control: Linear, equal percentage Recommended uses: 1. Fully open/closed or throttling services 2. Frequent operation 3. Minimal fluid trapping in line Applications: Liquids, gases, slurries, liquids with suspended solids Advantages High Capacity Good flow control Low cost and maintenance Low pressure drop
17.5
Disadvantages High torque required for control Prone to Cavitation at lower flows
OTHER VALVES 1) Another type of valve commonly used in conjunction with other valves is called a check valve. Check valves are designed to restrict the flow to one direction. If the flow reverses direction, the check valve closes. 2) Relief valves are used to regulate the operating pressure of incompressible flow. 3) Safety valves are used to release excess pressure in gases or compressible fluids.
17.6
CONTROL VALVES Control valves are of three types based on how the valve travel or stroke (openness) relates to the flow: • Equal Percentage: equal increments of valve travel produce an equal percentage in flow change • Linear: valve travel is directly proportional to the valve stoke • Quick opening: large increase in flow with a small change in valve stroke
17.7
SOME RULES OF THUMB Equal Percentage (most commonly used valve control) • Used in processes where large changes in pressure drop are expected • Used in processes where a small percentage of the total pressure drop is permitted by the valve • Used in temperature and pressure control loops Linear • Used in liquid level or flow loops • Used in systems where the pressure drop across the valve is expected to remain fairly constant (i.e. steady state systems) Quick Opening • Used for frequent on-off service • Used for processes where "instantly" large flow is needed (ie. safety systems or cooling water systems)
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN PROCESS DESIGN GUIDELINES
Rev.: A
Date: 08/02/08
Page 50 of 87
17.8
CONTROL VALVE FLOW COEFFICIENT
17.8.1 CV FOR INCOMPRESSIBLE FLUID The limiting pressure drop corresponding to the occurrence of the critical flow is defined as: ∆Pcritical = P1 − Pv Control valve flow coefficient can be calculated using the following formulas:
∆P = P1 − P2 Cf =
∆P ∆PVC NORMAL FLOW
CRITICAL FLOW
∆P < C f ⋅ ∆Pcritical
∆P ≥ C f ⋅ ∆Pcritical
2
C v = 1.16 × Q ⋅
FOR LIQUID
G ∆P
2
C v = 1.16 × Q ⋅
G C f ⋅ ∆Pcritical 2
Where P1 = Upstream pressure (bara) P2 = Downstream pressure (bara) Pv = Vapor pressure of liquid at flowing temperature (bara) Pc = Critical pressure of liquid (bara) Pvc = Differential pressure between the inlet pressure and the pressure at Vena Contracta (bar) T = Upstream flowing temperature (K) G = Specific gravity @STP Cf = Pressure recovery factor (assume 0.8 – 0.9 if no data is available) Cv = Valve flow coefficient Q = Volumetric flow rate (m³/h) @STP 17.8.2 CV FOR COMPRESSIBLE FLUID The limiting pressure drop corresponding to the occurrence of the critical flow is defined as: 2 ∆Pcritical = 0.5 ⋅ C f ⋅ P1 Control valve flow coefficient can be calculated using the following formulas:
∆P = P1 − P2 G=
ρg @ STP ρ AIR
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN
Rev.: A
PROCESS DESIGN GUIDELINES
Date: 08/02/08
Page 51 of 87
NORMAL FLOW
CRITICAL FLOW
∆P < 0.5 ⋅ C f ⋅ P1
∆P ≥ 0.5 ⋅ C f ⋅ P1
2
FOR GAS
Cv =
SATURATED STEAM
SUPERHEATED STEAM
Q G ⋅T 295 ∆P ⋅ ( P1 + P2 )
Cv =
Cv =
2
Cv =
72.4 × W
Cv =
∆P ⋅ ( P1 + P2 )
72.4 × W × (1 + 0.00126 ⋅ Ts ) ∆P ⋅ ( P1 + P2 )
Q ⋅ G ⋅T 257 ⋅ C f ⋅ P1
Cv =
83.7 × W C f × P1
83.7 × W × (1 + 0.00126 ⋅ Ts ) C f ⋅ P1
Where P1 = Upstream pressure (bara) P2 = Downstream pressure (bara) T = Upstream flowing temperature (K) Ts = Steam superheat temperature (°C) G = Gas specific gravity @STP Cf = Pressure recovery factor (assume 0.8 – 0.9 if no data is available) Cv = Valve flow coefficient W = Mass flow rate (tph) Q = Volumetric flow rate (m³/h) @STP
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN
Rev.: A
PROCESS DESIGN GUIDELINES
Date: 08/02/08
Page 52 of 87
17.9
CONTROL VALVE ISOLATION AND BYPASS VALVE SIZE CONTROL VALVE MANIFOLD
D
IF REQUIRED
D
3/4"
3/4"
1) For control valve diameter < 4" in continuous service, complete manifold (block+bypass) shall be provided 2) For control valve diameter > 4" the need for installation of block & bypass valve and handwheel around the control valve shall be defined for case by case 3) Bypass valves are globe valves for size 600# 2) Type of valve and isolation by spacer or spectacle blind must be as per the standard practice
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN
Rev.: A
PROCESS DESIGN GUIDELINES
Date: 08/02/08
Page 54 of 87
17.11 TYPE OF ISOLATION VALVE ACCORDING TO SERVICE Hydrocarbon Liquid/Gas
Φ ≤ 2” → Ball valve FB(
)
Valve upstream, downstream PSV → Ball valve FB ( ) BDV → Ball valve FB ( ) ESDV→ Ball valve FB ( )
Chemical Product, Air, Nitrogen, Inert gas & CO2 Φ ≤ 2” → Ball valve FB( )
Valve upstream, downstream PSV → Ball valve FB ( BDV → Ball Valve FB ( ) ESDV→ Ball valve FB ( )
Glycol Same criteria as hydrocarbon services if operating temperature is ≥ 170 oC, the type of ball valve will be revised.
Φ ≥ 2” → Butterfly valve Machinery suction: → Ball valve FB (
)
Φ > 2” → Butterfly valve for sea water.
)
Reduced ball valve ( ) will be used for the other cases. Also SDV (except if located at equipment suction).
Reduced ball valve ( ) will be used for the other cases. Also SDV (except if located at equipment suction).
The process will decide if ( ) is acceptable or not for the different services.
The process will decide if ( ) is acceptable or not for the different services.
Limit of ball valve utilization: T = 170 oC If operating temperature is ≥ 170°C, the type of ball valve will be revised.
Limit of ball valve utilization: T = 170 oC If operating temp. is ≥ 170°C, the type of ball valve will be revised.
INJECTION WATER P > 150 bar → Ball valve ( )
17.12 TYPCAL VALVE ARRANGEMENTS PRESSURISATION BY-PASS DOUBLE VALVING
3/4" 2" BYPASS
Φ < 2” → Gate valve
Φ ≤ 2” →Gate valve or diaphragm valve for sea water
)
Pressure valve to flare: → Ball valve FB ( ) Machinery suction: → Ball valve FB (
Sea Water, Cooling Water & Hot Water
D 2"
LC
BYPASS
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN
Rev.: A
PROCESS DESIGN GUIDELINES
Date: 08/02/08
Page 55 of 87
DETAILS OF SAMPLE CONNECTION AS IN PID DRAWING
SP
3/4"
Liquid sample connection with temperature < 65°C and a vapor pressure < ATM
300 min
1/2"
2"
DO DRAIN
UTILITY INJECTION Injection dia =3/4"
MIN CHEMICAL LINE
Sam e piping class Dia = 3/4" D
Top or Side connection
PRO CESS LINE
Rem ovable
UTILITY INJECTION Injection dia > 3/4"
UTILITY LINE
D
3/4"
FLOW Process Class
3/4"
Process or Utility class
Utility Class
PSV INSTALLATION TO FLARE HEADER
3/4"
LO
SP
3/4"
SPEC BREAK LO V
LO
3/4"
600 min
LO V
SPEC BREAK
LO V
PSV
SPEC BREAK
PSV
SPEC BREAK
PSV
LC
NOTES: 1) Distance between PSV and reducer is minimum 2) Downstream block valve is installed flange to flange (if the builtup back pressure is high this valve is installed downstream the reducer) 3) The minimum distance of 600 mm is required in case of cold depressurisation to avoid icing of the ball valve
NUMBER OF PSV & SET PRESSURE 1) Only one PSV is installed for Thermal or Fire relief case 2) 1+1 or multiple PSVs are installed for other relief cases. The total relief load is shared by all the operating PSVs 3) For multiple PSVs one PSV is set at Design pressure and is in operation while a similar PSV is in standby. All other PSVs are set at 105% of set pressure and are in operation.
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN
Rev.: A
PROCESS DESIGN GUIDELINES
Date: 08/02/08
Page 56 of 87
PUMPS HANDLING HYDROCARBON & GLYCOL
UC 3/4" UC 3/4"
3/4"
3/4"
3/4"
NOTES: 1) "Y" type strainer are installed for lines 80°C for spare pump heating or for draining the spare pump discharge side 4) Utility connection (UC) is required only for hydrocarbon services
COLLECTED DRAIN
DRAINAGE PIPING FROM HORIZONTAL VESSELS
Level instr. drain
DRAINAGE PIPING FROM VERTICAL VESSELS AND PROCESS PIPING
PROCESS LINE Level instr. drain
COLLECTED DRAIN
COLLECTED DRAIN
2" D
2"
3/4" SPEC BREAK
D
3/4" SPEC BREAK
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN PROCESS DESIGN GUIDELINES
Rev.: A
Date: 08/02/08
Page 57 of 87
18.
FLARE SYSTEM Generally there are two types of flares: • Ground flare (in case of more land area available) • Elevated flare (common practice) Burning pit, an excavated pit of 5 ft depth, with a burner inserted in a wall. These are not generally used because of excessive smoke. Location - at least 500 ft from process and storage area and from any roadway. Flares can be designed for smokeless operation. Two types of smokeless flares are: • Steam sparged flare • Air assisted flare The flare system shall be any one or combination of the following: • HP flare system • LP flare system The general philosophy for LP and HP flare system in an Oil & Gas facility are: 1) Relieving devices with set pressure 10 barg are connected to the HP flare system. Relief devices can not relieve the design relief rate, when the back pressure is more than certain limit, as stated below: 1) Back pressure for conventional safety valves shall be < 10% of set pressure. 2) Back pressure for balanced bellows safety valves shall be < 50% of set pressure. 3) Back pressure for pilot operated safety valves shall be < 90% of set pressure.
18.1
ELEVATED FLARE Elevated flares consists of a stack (guyed, self-supporting or with a supporting structure) with a burner tip, pilot burners and associated fuel system, igniter and miscellaneous auxiliaries. Recommended distances of the flare stack from other equipments are: 1) >200 ft (61 m) to a separator/floating roof tank which could be ignited by an occasional falling spark 2) > 400 ft (122 m) to any other equipment 3) > 500 ft (152 m) to any equipment having an elevation within 125-150 ft of the elevation of the flare tip Flare stack diameter is generally sized on a velocity basis, although pressure drop should be checked. Diameter of the flare stack is to be not less than that of the flare tip. Flare tip shall be sized for Mach 0.2 to 0.4 at max emergency load. The higher limit is to be used with fuels with high burning velocities. For example, small amounts of H2 in a hydrocarbon fuel increases the burning velocity and permits higher gas velocities in the tip without flame blow off. Minimum velocity at full capacity should be 100 fps. Too low a tip velocity can actually cause heat and corrosion damage. Flare stack height is generally designed based on the radiant heat intensity generated by the flame. A check should also be made on the max ground concentration level if toxic and corrosive pollutants are present in the stream.
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN PROCESS DESIGN GUIDELINES
Rev.: A
Date: 08/02/08
Page 58 of 87
Some heat intensity levels which have been employed in design are: 1) 3.16 kW/m² (1000 Btu/h ft²): storage tanks containing volatile materials and control rooms or areas where personnel must remain at their posts. 2) 4.73 kW/m² (1500 Btu/h ft²): used by some companies for acceptable intensity in operating areas where operators wearing normal clothing were likely to perform their duties and where general area radiant shielding exists. 3) 6.31 kW/m² (2000 Btu/h ft²): in open areas where no shelter is available and only escape is required. 4) 15.77 kW/m² (5000 Btu/h ft²): on structures and in operating areas where operators are not likely to perform their duties and where shelter from radiant heat is available, e.g. behind equipment. Elevated Flare Flash Back Protection Protection shall be provided against flash back from the flare tip to the flare header and blowdown drum. For this purpose gas purging in connection with sealing devices are recommended. Two methods of sealing are : (a) Molecular seal, (b) Liquid seal Liquid Seal A liquid seal is often provided in the flare stack either at the base of the stack itself or in a separate seal drum. The preferred liquid is water although oil may also be used. In sizing a seal drum, first determine the maximum pressure on the flare side. The minimum seal length provided by the water discharge pipe is then specified for 200% of maximum operating pressure of the drum or 10 ft, whichever is greater. The drum diameter is set to provide sufficient vapor/liquid disengaging area, with the vapor velocity V limited to Vc. (D − d ) V = 0.157 d V = velocity (ft/s) D = density of entrained liquid, (lb/ft³)
d = density of gas (lb/ft³)
The height of vapor space in a vertical drum should be approx 2 - 3 times the drum diameter. If a horizontal vessel is used a minimum dia of 3 ft is recommended for this purpose. A seal drum design pressure of 50 psig is suggested to protect the drum from pressure surges, as with the blowdown drum. However if continuous gas purging is not used, the seal drum design pressure should be 150 psig in order to withstand the overpressure due to explosion. Gunite lining is normally used for corrosion control. Pipe Seal Pipe seal usually consisting of a loop or trap built into the base of the flare stack or the flare line inlet connection, are designed to be filled with water to prevent flashback into the flare headers. The slop of the inlet is designed to provide a volume of water below the normal sealing water level equivalent to the volume of 10 ft of the inlet line. Although the 10 ft volume basis is arbitrary, it should provide adequate sealing against flashback. Depth of the water seal is usually held to about 12 in since greater depths can cause gas pulsations. Seal water level is maintained by a continuous flow of water at about 20 gpm. Smokeless Operation A diluents such as steam, air or water may be used to improve flame characteristics i.e. reduce smoke and luminosity. Steam is preferred to compressed air because of cost. Steam can be injected into the flare stack below the burner tip. Steam required for smokeless operation can
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN
Rev.: A
PROCESS DESIGN GUIDELINES
Date: 08/02/08
Page 59 of 87
be calculated from the following formula : 10.8 ⎞ ⎛ WS = WHC ⎜ 0.68 − ⎟ M ⎠ ⎝ Ws = Steam rate (lb/h) Whc = hydrocarbon rate (lb/h)
M = Gas MW
When unsaturated hydrocarbons are present (which are more difficult to burn) following steam rates shall be used :
Steam rate (kg steam / kg gas)
0.25
For no spoke operation ---
For reduced luminosity ---
10
0.35
---
---
20
0.6
---
---
30
0.75
---
---
35
0.88
1.55
---
40
1
1.75
2.75
50
1.2
2.15
3.2
60
1.45
2.55
3.65
65
1.55
2.75
3.85
% unsaturates
No smoke operation
5
Smokeless operation is usually specified for only 10 - 30% of the max flare load. Ignition The number of pilot units (with remote ignition) required per flare is largely a function of wind conditions. If the prevailing wind is strong, then one pilot unit may be adequate. Standard flare tip includes three pilot units uniformly distributed around the top of the flare. Smokeless operation 1) The stack itself will be of carbon steel. The foot of the stack (below the main seal) is cement filled 2) The bottom of the stack and inlet are to be gunite-lined from the cement fill to a height of 6'0" above the sealing water level. Dip leg and bottom of flare stack are to be insulated and steam traced to a height of 6'-0" above sealing water level (to vaporize condensed hydrocarbons) 3) The top 10' of flare stack and pilot assembly to be type SS-310. If steam is injected in to the flare stack, below the burner tip, the entire tip assembly will be of heat-resisting alloy 4) Sometimes, when a gas stream has a continuous high level of H2S, there is an economic advantage to providing a separate flare, header and seal drum for the H2S rich stream Noise The roar of combustion is the one source of noise that cannot be avoided with elevated flares. At moderate release rates this source generally does not cause problem, but if the roar remains objectionable, the moderate (and most frequent) release can be burned in a ground flare, keeping the elevated flare for emergencies. In smokeless flares, generally steam is used which increases the combustion roar. In this case the noise can be reduced to some extent by using a multipoint steam nozzle.
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN PROCESS DESIGN GUIDELINES
Rev.: A
Date: 08/02/08
Page 60 of 87
18.2
GROUND FLARE Ground flare is similar to an elevated flare, but without any supporting structure. It requires 76 – 152 m (250 - 500 ft) sterile distance. A ground flare may be either single tip or multi-jet type. The former is no longer recommended. All new installations are multi-jet type. A multi-jet ground flare is designed for burning gas at nearly atmospheric pressure. The gas is supplied by a manifold of parallel burner lines which lead to a number of vertically discharging 1" diameter burners. Solid rods acting as flame holders are mounted above the outlet of the burners parallel to the burner lines, their main function being to prevent the flame from riding up to the top of flare. The flames are surrounded by a refractory lined stack. The stack acts as a chimney, drawing air in to the combustion zone. The stack in turn is surrounded by a windbreak to reduce wind effects and luminosity at ground level. The use of a ground flare requires a clear distance of about 500 ft from any processing unit or storage facility and not less than 200 ft from property line fence. Design Criteria Large flow rates cannot be treated by a ground flare; John Zinc suggests a max. flow rate of 40000 - 60000 lb/h. To determine the diameter of a ground flare, the ESSO formula for multijet flares can be used : D = 0.826 Q D = Stack dia (ft)
Q = Heat release (MMBtu/h)
For stack diameters up to 25 ft, the stack length is usually made 32 ft. For larger diameters, the stack height should be suitable increased. The bottom of the stack is elevated to allow air for combustion to enter. Minimum clearance between bottom of stack and grade is either 6 ft or 0.3 D, whichever is greater. Sealing Device The same for elevated flare can be used. A special sealing device "two stage water seal for use with integrated flare system" has also been developed for a combined elevated and ground flare. The base of the elevated flare contains a seal drum with two dip legs. Gas is first sent through the lower seal to ground flare. Up to the capacity of the ground flare, all gas is sent to the ground flare, with increasing gas flows handled by varying the number of burners operating. As the gas flow increases further the upper seal is broken and gas is sent to the elevated flare. Smokeless operation Smokeless flaring - both ground & elevated flares can be operated under smokeless condition by suitable injection of steam or air. However smokeless firing under all conditions is uneconomic and smokeless operation is generally specified only for 10-30% of the maximum flare load. Smoke production, depending upon MW and gas flow can be reduced in multi-jet ground flares with steam or water injection. However a multi-jet ground flare is less smoky than an elevated flare without steam injection. Material of construction Ground flares are generally constructed of fire brick or of a CS cylinder lined with about 8" of refractory material.
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN PROCESS DESIGN GUIDELINES
Rev.: A
Date: 08/02/08
Page 61 of 87
Noise From available data and consequent extrapolations, it has been found that, in the absence of combustion- drive pulsation, the grade level flare is likely to be about 10 dB quitter at the same load than an elevated flare. 18.3
VENT STACK If vapors from pressure relief valves may be directly discharged to atm, this can be done by simply piping to a safe location (called as vent pipe) or by collecting these vapors via laterals and headers and disposing of them together with other valve emissions to the atmosphere through a common stack (called as vent stack). Air dilution The mass flux (in lb/h) when discharged directly to the atm may be expressed by the following equation: W = 0.264 ⋅ X / D Wo Where:X = distance from the jet nozzle (vent pipe end) (ft) D = vent pipe (nozzle dia) (ft) W = weight flow rate of vapor-air mixture at the distance X from the vent pipe end (lb/h) Wo = weight flow rate of the vapors discharged by the vent pipe (lb/h) Based on this info, hydrocarbon vapors discharging from a safety valve into the atmosphere will entrain sufficient air to be below their LEL at a distance approx 120 tail pipe diameters away from the end of the vent. In order to supply the energy requirement for this amount of mixing, it is necessary that the velocity of the discharged hydrocarbon vapors be more than 152 m/s as they leave the tail-pipe. When reduction of pipe size is required to maintain 152 m/s or higher outlet velocity, the pipe end should be reduced for a length equal to 3 pipe diameters. Location of vent stacks 1) At least 1.83 m (6 ft) above the highest adjacent structure or tower. 2) At least 3.66 m (12 ft) above the highest adjacent platform. 3) At least 4.57 m (15 ft) above grade. 4) At least 15.2 m (50 ft) or 120 pipe diameters, whichever is greater, away from the nearest platform structure or tower when located at an elevation lower than the platform, structure or tower. 5) At least 30.5 m (100 ft) or 120 pipe diameters, whichever is greater, away from the tops of flue gas stacks or other ignition sources regardless of the atmospheric vent elevation. Design Criteria High discharge velocities are desirable in order to obtain good hydrocarbon-air mixing and reach a concentration below 3% by weight within a reasonable distance from the stack. Tip Diameter The tip dia. is usually sized for an exit velocity of 500 fps at the max relief rate. Sonic velocity is to be avoided.
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN PROCESS DESIGN GUIDELINES
Rev.: A
Date: 08/02/08
Page 62 of 87
π ⋅ DO 2 4
=
WMAX 3600 ⋅ d V ⋅ U O ⋅ MAX
Where Do = Stack tip outlet inside dia (ft) Wmax = Max relief rate (lb/h) dv = Vapor density (lb/ft³) Uo max = Max allow exit velocity (ft/s) (500 fps or about 0.5 sonic velocity, whichever is lower) Stack Diameter Not less than tip diameter. Total pressure drop due to header, stack, tip and exit loss shall be checked against available pressure drop. Stack Height Vent stack height shall be selected to ensure that : 1) Toxic compound concentration at ground or at any other point of interest is lower than the threshold limit. If the pollutant is odorous, the odor threshold of perception at ground level inside / outside the plant should be checked. 2) Radiant heat intensity, in the event the relieved stream should become ignited, does not exceed the limits valid for flare stacks. 3) Vapor concentration at ground or at any other point of interest is limited to 0.1-0.5 times the lower flammability limit of the vapor Silencing Silencers are sometimes provided in vent stacks to reduce noise caused by high velocity gas discharge into atmosphere. Silencers are generally provided in case of continuously operated vent stacks. Sealing In every vent stack installation, careful consideration should be given to the problem of any accumulation of liquid in the line leading to the stack. To cope with this situation, it is important to avoid any pocket in the line and to slope the system to a low point drain. Seals shall be provided with a ht. equiv. to at least 1.5 times the back pressure under max relief load. Snuffing Steam The possibility that the vent stack may become accidentally ignited by lightning or from other sources usually makes a remotely controlled snuffing steam connection desirable on a vent stack or any atm vent handling flammable vapors. It is impractical to size the steam supply line for a rate sufficient to extinguish a fire under max venting conditions. Purging Continuous purging of vent systems should be provided in case discharge of flammable vapors or gases is anticipated through long headers leading to the final vent stack. (purge rate shall be as given in elevated flare)
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN PROCESS DESIGN GUIDELINES
Rev.: A
Date: 08/02/08
Page 63 of 87
18.4
FLARE HEADER AND SUB-HEADER SIZING Network Sizing Criteria
Fluid Phase
Gas Phase Two Phase
TOTAL Spec GS ECP 103
Service
SAIPEM Standard
Mach No.
ρV² (kg/m s²)
Mach No.
ρV² (kg/m s²)
Lines downstream of relieving device and sub-headers
0.7
< 150 000
0.5
< 100 000
Headers
0.7
< 150 000
0.7
< 100 000
Lines downstream of relieving device, sub-headers and headers
0.25
< 50 000
Lapple method and charts For vapors and gases Lappel method may be used for both adiabatic and isothermal conditions. Although flow in a relief header approaches adiabatic conditions, no significant errors are made if calculations are limited to isothermal flow because hydrocarbons above 50 MW have a specific heat ratio, k, approaching 1. This corresponds to isothermal expansion. The slight errors deriving from the assumption of isothermal flow are on the conservative side. Lappel chart for isothermal flow to be used. Corresponding charts for adiabatic flow can be found in Perry handbook. Calculation
Gci = 12.6 × Po
M ((2 ⋅ Z ) − 1) × To
Where Gci = critical mass flow (lb/s ft²) Po = upstream pressure (psia) To = upstream temperature (R) M = MW of vapor Z = compressibility factor, at flow conditions or average over the pipe length
N = 4× f ×
Le D
Where N = line resistance factor D = ID of pipe (ft) Le = equivalent length of pipe & fittings (ft) f = Fanning friction factor; initial estimate of 0.004 is suggested. This can be rechecked later using Reynolds number. Note When Darcy or Moody friction factor is available, their value must be divided by 4, to obtain Fanning friction factor. When analyzing a relief header, made up of various sections of different diameter, each
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN PROCESS DESIGN GUIDELINES
Rev.: A
Date: 08/02/08
Page 64 of 87
section with dia D and resistance factor N can be converted to an equivalent diameter De having a line resistance factor Ne equal to
⎛D ⎞ Ne = N × ⎜ e ⎟ ⎝ D⎠
4
Any diameter may be selected as equivalent diameter (De), but the calculation of mass velocity (G and Gci) must be referred to this diameter De. The total resistance factor Nt of the entire header, referred to the equivalent diameter, De is the sum of equivalent resistance factors Ne of the various sections of the header. The pressure drop of the header can be found using Lappel chart with Nt and G/Gci calculated with equivalent diameter De. Compressible fluids & flashing liquids For the sizing of a blow down header carrying saturated liquids flashing as the pressure drops, the following equation may be used: (N × Va / 2) + ∆V ∆P = P1 − P2 = 7.26 ⋅ w 2 ⋅ D4 Where P1 = inlet pressure (psia) P2 = outlet pressure (psia) w = mass flow rate (liquid + gas) (lb/s) N = number of velocity heads, as defined in Lapple method Va = average specific volume (ft³/lb) D = pipe diameter (inch) ∆V = V2 – V1 = increase in specific volume of fluid between inlet and outlet, (ft³/lb) Suggested velocities and sonic flow Sonic velocity may be calculated from the following formula (valid for ideal gas):
VS = 223
(Z ⋅ k ⋅ T )
M Where Vs = sonic velocity (ft/s) K = Cp/Cv at fluid temperature MW = molecular weight T = fluid temperature (°R) Z = compressibility factor (approx. 1 at low pressure) Even if sufficient pressure is available, depending on the lowest set pressure of relief valve in the system, it is not desirable to size the header so that the flow becomes sonic (high noise level and pipe vibrations). To prevent this, the value for the ratio (P2/Po) / (G/Gci) must be larger than 0.6 and preferable not less than 1, P2 being the pressure in the downstream vessel. The recommended range of values for flare headers and other piping when pressure drop is not controlling are listed below:
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN
Rev.: A
PROCESS DESIGN GUIDELINES
Date: 08/02/08
Page 65 of 87
Flare System
18.5
(P2/Po) / (G/Gci)
V / Vs
Downstream piping of safety valves
1 to 2
0.25 to 0.65
Flare headers
1 to 2
0.25 to 0.65
Flare stacks
2 to 3
0.20 to 0.40
FLARE KO DRUM SIZING A knock out (KO) drum is provided in the flare system to collect the liquid in the flare gas before being set to the flare stack. Flare KO drum shall be sized according to API 521. Flare KO drum shall be sized to meet the following criteria: 1) KO drum shall be sized to hold the maximum liquid relief in any relief case or scenario, between normal liquid level (NLL) and high liquid level (HLL or LHH), for a relief duration of 5 minutes (liquid hold-up time for general oil & gas facility). This hold-up time varies for each type of project. 2) The vapor space above high liquid level or high high liquid level (LSH, LAH, LSHH, LAHH) shall be sufficient to allow the maximum gas relief rate with a particle size separation of >600 microns (as per API 521) or >400 microns (as per TOTAL spec). 3) KO drum type shall be selected based on the following considerations: 4) If there is no space constraint, generally horizontal KO drum is preferred. 5) The residence time for liquid separation, is more in horizontal KO drum than vertical drum. If the liquid relief rate is more to the flare system, horizontal KO drum shall be selected. 6) The available liquid holdup volume is more in a horizontal drum than in a vertical drum for the same amount of gas rate. Therefore, horizontal drum shall be selected for system having substantial liquid relief. 7) Vertical KO drums are preferred due to tight space constraints. Vertical KO drums are selected for gaseous relief systems, where no or less liquid carryover is expected. 8) The suggested design pressure of blowdown drum is 50 psig to offer a satisfactory resistance in case of overpressure 9) A steam coil is generally provided in the blowdown drum to prevent heavy hydrocarbons from reaching their pour point. For the same reason, tracing is normally provided on suction and discharge lines of the blowdown pumps 10) The blowdown pumps can be motor or steam turbine driven. It is advisable for safety to have one pump motor driven and one steam driven 11) Due to variety of liquids handled, SG value is usually given as a wide range (e.g. 0.6-0.85) 12) A water boot may or may not be provided below the drum 13) Blowdown drum should be sloped 1:250 towards the liquid outlet nozzle 14) Horizontal KO drums shall be provided with either single pass or two pass arrangement to optimize the design. The different types of KO drum are shown below:
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN PROCESS DESIGN GUIDELINES
Rev.: A
Date: 08/02/08
Page 66 of 87
Flare Gas
Relief Gas
Flare Gas
Relief Gas
FLARE KO DRUM
FLARE KO DRUM (Single pass)
Flare Gas Relief Gas
Flare Gas Relief Gas
FLARE KO DRUM (Tw o pass)
FLARE KO DRUM (Tw o pass)
15) For most of the oil & gas facility, high high level in flare KO drum initiates plant shutdown system. Therefore, KO drum sizing should be done carefully to hold the emergency liquid relief. KO drums are provided with the following equipment / accessories: 16) KO drum pump to evacuate the liquid for further processing. The KO drum pump shall be sized to evacuate the liquid collected during emergency relief in 20 – 30 minutes. 17) KO drums are sometimes provided with boot to enhance the suction head to avoid NPSH problem for the KO drum pump. The reason for providing boot are: • Flare headers are routed with slop towards the KO drum, and hence the KO drum is located at ground level or lowest level in the plant. There would not be sufficient liquid head for the pump for safe operation. To avoid cavitation in the pump and to boost the NPSH, boots are provided from which the pump suction is taken. • The liquid collected in the KO drum are mostly at bubble point or flashing nature. This creates cavitation problem in the pump. By adding a boot of sufficient height, the suction static head and NPSH can be improved. • For satisfactory performance of any pump, the available NPSH should be >1 meter. In cases, where the NPSH is lower than 1, which is a typical case for KO drum pump, special pumps like canned pumps are selected. 18) KO drums are also provided with internal / external heaters. The reasons for the heater are: • To keep the liquid warm, to avoid freezing of the liquid or to reduce the viscosity of the liquid. • Internal heaters are sometimes used to vaporize the light liquid collected in the KO drum, so that they can be flared. For such cases pumps may not be provided for the KO drum. • Generally electrical heaters are used in flare KO drum due to the remote location in the plant and their rare operational requirement.
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN PROCESS DESIGN GUIDELINES
Rev.: A
Date: 08/02/08
Page 67 of 87
19) Following procedure from API 521 shall be used for flare KO drum sizing: Liquid particles will separate when: • the residence time of the vapour or gas is equal to or greater than the time required to travel the available vertical height at the dropout velocity of the liquid particles • the vertical gas velocity is sufficiently low to permit the liquid dropout to fall The dropout velocity of a particle in a stream is calculated using the following equation:
U C = 1.15 × g ⋅ D ⋅ Where Uc g D ρl ρv C µ
(ρ l − ρ v ) ρv ⋅ C
= dropout velocity (m/s) = acceleration due to gravity (9.81 m/s²) = particle diameter (m) = density of the liquid at operating condition (kg/m³) = density of the vapour at operating condition (kg/m³) = drag coefficient = viscosity of the gas (cP)
The drag coefficient (C) is calculated using the following equation: C ⋅ (Re) 2 = 0.13 × 10 8
ρ v ⋅ D 3 (ρ l − ρ v ) µ2
[ (
C = 10^ (- 0.0001 Log C ⋅ Re 2
)]
4
[ (
- 0.0028 Log C ⋅ Re 2
)]
3
[ (
+ 0.1225 Log C ⋅ Re 2
)]
2
[ (
- 1.136 Log C ⋅ Re 2
(The drag coefficient (C) can also be taken from Figure-20 on page 64 in API 521) The liquid dropout time is calculated as:
θ = hv / U c Where θ = dropout time (sec) hv = vertical height available above liquid level in KO drum (m) The actual vapour velocity is calculated as:
Uv =
Qv / 3600 ( Av ⋅ N b )
Where Qg = volumetric vapour flow rate (m³/h) Av = vapour cross sectional area (m²) Nb = number of vapour passes The required drum length is calculated as:
)]
+ 2.8091)
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN PROCESS DESIGN GUIDELINES
Rev.: A
Date: 08/02/08
Page 68 of 87
Lmin = U v ⋅ θ ⋅ N b The selected KO drum length should be > Lmin (Otherwise the above calculation to be repeated until the required KO drum length has been arrived) According to in-house KO sizing criteria, the gas flow area of KO drum shall meet one of the following requirements of λmax:
λ max =
Q *max Av⋅⋅ min
Q *max = Qv⋅⋅ max ⋅
ρv ρl − ρv
Where λmax = Gas load factor (m/s) = 0.07 m/s for general sizing = 0.10 m/s for KO drum with no inlet device and significant liquid quantities = 0.15 m/s for KO drum with Schoepentoeter inlet device and significant liquid quantities = 0.25 m/s for KO drum with relatively dry gas feed and/or with battery limit KO drums (generally selected) Qv, max = Maximum gas flow rate (m³/s) ρv = Density of gas (kg/m³) ρl = Density of liquid (kg/m³) Av, min = Minimum vapour flow area (m²) 18.6
TYPES OF FLARE HEADERS At times it may be economical to provide more than one header for the following reasons: 1) A small header is provided to handle relief streams which require sophisticated materials of construction for corrosion or design temperature reasons, while a large diameter header handles the remaining relief streams. 2) Two separate headers might prove economic when a considerable quantity of gas to be relieved comes from relief valves set at high pressure. In this case a high back pressure can be chosen to design one of the headers, using high velocity and smaller diameter, while other discharges are sent to a low pressure header. Header pressure drop, DP ~= W² / D W = mass flow rate, kg/h D = density, kg/m³ For flare load, at a time single risk is assumed. For fire case a fire area of 185 - 465 m² (2000 5000 ft²) can be assumed. Generally, 232 m² (2500 ft²) of fire area shall be taken to calculate the relief load. Areas larger than 465 m² (5000 ft²) are to be considered if enclosed by a curb or dike. Fire areas should be smallest areas which can be isolated by means of fire fighting equipment. In case of power failure, air coolers still absorb 20-30% of design heat duty by natural convection.
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN PROCESS DESIGN GUIDELINES
Rev.: A
Date: 08/02/08
Page 69 of 87
Relief of acid or toxic gases Hazardous gases such as H2S, Cl2, HCN, CO, etc. should be processed in order to convert them to a less objectionable compound; e.g. H2S to liquid sulfur, Cl2 to NaOCl with NaOH wash, etc. Alternately it might be desirable to incinerate these gases in a properly designed incinerator. However during emergencies, it may be necessary to discharge these gases to flare. In this case, special precautions must be taken, such as:• A separate blow down header, without pockets, with suitable materials to prevent corrosion. In the case of wet H2S it is desirable to insulate and steam trace the line. • The hazardous gas, piped thro a separate header, can be discharged either to the common flare (downstream the seal or at flare tip) or preferable, to a separate flare for better burner control. 18.7
FLARE HEADER PURGING Explosive conditions can occur if air enters the relief header or flare system. Air may enter the system when (1) the gas density is lower than air, (2) air can enter the flare stack while gas escapes, (3) at the end of hot gas flaring, during system cooing and creation of vacuum causes air to be sucked into the system, (4) natural draft of flare headers may cause, in case of leaking joints, air to enter. Water seal and molecular seals can be provided to prevent air diffusion in to the flare header. In freezing climates, steam injection and heating coils to be provided for water seals. It is also necessary to provide the injection of gas (fuel gas or inert gas) for the following reasons: • To provide a continuous purge of the system • To prevent vacuum condition when the system cools down after hot gas flaring. Purge gas velocity in the header can be 0.61 - 0.91 m/s (2-3 ft/s), for flares without water seal and molecular seal. With sealing devices, a purge rate of 0.015 - 0.091 m/s (0.05-0.3 ft/s) is generally satisfactory. Higher purge velocity is required for low molecular weight gas and higher stack diameters. Safe condition exists, if O2 concentration is kept lower than 50% of the LEL inside the vent stack, at 7.62 m (25 ft) from the top. Purge gas velocity is a function of diameter of stack and MW of the purge gas. The table below gives typical purge gas rates for various gas MW and stack diameters. The purge gas rates given are for vent stacks. For flare stacks, lower purge gas rates (50% of given values) can be used because the presence of the flame is a deterrent to oxygen diffusion. Recommended purge gas velocity (m/s) Molecular weight of gas 8 12 20 Dia of stack 10" 0.006 0.005 0.002 20" 0.018 0.012 0.009 30" 0.052 0.043 0.037 40" 0.168 0.015 0.137
28 0.002 0.006 0.034 0.128
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN PROCESS DESIGN GUIDELINES
Rev.: A
Date: 08/02/08
Page 70 of 87
Purge gas rate required to prevent vacuum conditions after a hot gas flaring, can be calculated by the following formula :
G = 60 × VO⋅ ×
(1 − (Ta / T )) t
Where G = Purge gas flow rate (ft³/s) Vo = Total flare system volume (ft³) T = Hot gas temperature (R) Ta = Ambient air temperature (R) t = Time for cooling from T to Ta (min) (~ 10-20 min)
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN PROCESS DESIGN GUIDELINES
Rev.: A
Date: 08/02/08
Page 71 of 87
19.
HEAT TRANSFER FLUID (HTF) SYSTEM DESIGN HOT OIL SYSTEM CONFIGURATION
CONSUMERS N2 HOT OIL MAKEUP
TC PC
TRIM COOLER
SURGE TANK
FC
HOT OIL SUPPLY PUMP
TC
FIRED HEATER
Surge vessel is sized based on the following criteria: 1) To provide adequate volume to accommodate the fluid thermal expansion when heated from ambient to normal working temperature 2) To provide adequate NPSH for the HTF pumps 3) The surge drum is located at an elevation such that the normal operating HTF level in the drum will be located above the elevation of the highest component in the HTF system The volume between the minimum and normal working level shall be at least the larger of the following: 1) The volume increase of the total HTF inventory when the temperature is raised from minimum operational to normal working level 2) The volume of HTF lost via a ruptured tube in 15 minutes in a heat consumer operating with a process pressure below the HTF system pressure. The volume between the normal and maximum working level shall be at least : 1) The volume increase of the total HTF inventory when the temperature is raised from normal to maximum working level 2) Spill-over control valve and piping shall be designed for 30% of pump capacity 3) Fluid velocity shall not exceed 2 m/s 4) U-tube exchanger shall be used if hot oil is used in tube side 5) Moderate operating temperature of hot oil system is 275°C 6) Fouling resistance of 0.00017 m²K/W shall be used for HTF service Heat transfer fluids used are: petroleum oils below 600°F (315°C), Dowtherms or other Synthetics below 750°F (400°C), molten salts below 1100°F (600°C)
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN
Rev.: A
PROCESS DESIGN GUIDELINES
Date: 08/02/08
Page 72 of 87
20.
COOLING WATER SYSTEM DESIGN 1) Water in contact with air under adiabatic conditions eventually cools to the wet bulb temperature 2) Tower fill is of a highly open structure so as to minimize pressure drop, which is in standard practice a maximum of 2” of water 3) Chimney-assisted natural draft towers are of hyper-boloidal shapes because they have greater strength for a given thickness; a tower 250 ft high has concrete walls 5”-6” thk. 4) With industrial cooling towers, cooling to 90% of the ambient air saturation level is possible 5) Relative cooling tower size is dependent on the water temperature approach to the wet bulb temperature: Twater out - Twb
Relative Size
5
2.4
15
1.0
25
0.55
COOLING WATER SYSTEM CONFIGURATION
Soft Water Makeup
H2SO4 DOSING TANK
CHEMICAL DOSING TANK
CONSUMERS ID FAN
CELL-1
CELL-2
CELL-3
COOLING TOWER BASIN
SCREENS BLOWDOWN
EMERGENCY CW PPUMPS
COOLING WATER CIRCULATION PUMPS
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN PROCESS DESIGN GUIDELINES
Rev.: A
Date: 08/02/08
Page 73 of 87
General consideration 1) Water circulation rates are generally 2-4 gpm/ft² (81-162 lit / min / m²) and air velocities are usually (1.5-2.0 m/s) 2) Countercurrent induced draft towers are the most common. These towers are capable of cooling to within 2°F (1.1°C) of the wet bulb temperature. Approach of 5-10°F (2.8-5.5°C) is more common 3) Evaporation losses are about 1% by mass of the circulation rate for every 10°F (5.5°C) of cooling. Drift losses are around 0.25% of the circulation rate. A blow down of about 3% of the circulation rate is needed to prevent salt and chemical treatment buildup 4) Cooling tower water is received from the tower between 80-90 °F (27-32 °C) and should be returned between 115-125 °F (45-52 °C) depending on the size of the tower. Seawater should be return no higher than 110 °F (43 °C) 5) Cooling water systems shall be designed for the conditions specified in the data / requisition sheet, and in no case shall be designed for a working pressure of less than 7 barg 6) Provision shall be made for complete venting and draining of the system.
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN PROCESS DESIGN GUIDELINES
Rev.: A
Date: 08/02/08
Page 74 of 87
21.
REFRIGERATION SYSTEMS 1) A ton of refrigeration equals the removal of 12,000 Btu/h (12,700 kJ/h) of heat 2) For various refrigeration temperatures, the following are common refrigerants: Temp (°F)
Temp (°C)
Refrigerant
0 to 50
-18 to -10
Chilled Brine or Glycol
-50 to 0
-45 to -18
Ammonia, Freon, Butane
-150 to -50
-100 to -45
Ethane, Propane
3) Compression refrigeration with 100°F condenser requires the following power @HP/ton at various temperature levels: 1.24 at -6.7°C; 1.75 at -17.8°C; 3.1 at -40°C; 5.2 at -62.2°C 4) Below -62.2°C, cascades of two or three refrigerants are used 5) In single stage compression, the compression ratio is limited to about 4 6) In multistage compression, economy is improved with inter-stage flashing and recycling, so-called economizer operation 7) Absorption refrigeration (ammonia to -34.4°C, lithium bromide to +7.2°C) is economical when waste steam is available at 12 psig or so
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN PROCESS DESIGN GUIDELINES
Rev.: A
Date: 08/02/08
Page 75 of 87
22.
CHILLED WATER SYSTEM CHILLED WATER SYSTEM CONFIGURATION SOFT WATER or DM WATER MAKEUP
COOLING WATER HOT WELL
CHILLER UNIT
CONSUMERS
COLD WELL =30%-40% of vessel diam eter
CHILLED WATER TANK
BLOWDOWN
CHILLED WATER PUMPS
CHILLED WATER CIRCULATION PUMPS
General consideration 1) DM water is used as make-up water for the chilled water system 2) Chilled water supply temperature is approx 5°C to 10°C. Return temperature is 10°C to 15°C 3) The tank / pump / piping material of construction is Carbon Steel
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN
Rev.: A
PROCESS DESIGN GUIDELINES
Date: 08/02/08
Page 76 of 87
23.
CHILLED BRINE SYSTEM
CHILLED BRINE SYSTEM CONFIGURATION BRINE MAKEUP
CONSUMERS
CHILLED BRINE TANK
COOLING WATER
CHILLER UNIT
DRAIN
CHILLED BRINE CIRCULATION PUMPS
General consideration 1) 40% ethylene glycol solution is used as the brine solution 2) Chilled brine supply temperature is approx -15°C to -5°C. Return temperature is -10°C to 0°C 3) The tank / pump / piping material of construction is Carbon Steel and some times LTCS
Doc. N°: WIP-SIPS-PCS-001
WORK IMPROVEMENT PLAN
Rev.: A
PROCESS DESIGN GUIDELINES
Date: 08/02/08
Page 77 of 87
24.
DM WATER SYSTEM DM water is used as make up water for boiler feed water to generate steam. DM water is some times used as Process water and diluents for acids and chemicals preparation. DM WATER SYSTEM CONFIGURATION ACTIVE CARBON FILTER
SAND FILTER
SOFTENER
CATION EXCHANGER
WEAK BASE ANION EXCHANGER
RAW WATER
BLOWER DEGASSER
BACK WASH WATER HCl TANK
NaOH TANK DEGASSER PUMP
DRINKING WATER
SOFT WATER FOR COOLING WATER MAKEUP
MIXED BED EXCHANGER
STRONG BASE ANION EXCHANGER
DM WATER PACKAGE
DM WATER TANK
DM WATER CONSUMERS DM WATER PUMPS
HCl TANK
ACID DRAIN
NaOH TANK
Boiler Feed Water (BFW) specification is presented in the following table: Sl.No. 1. 2. 3. 4. 5. 6. 7. 8.
Characteristics Total hardness as CaCO3 max. (mg/lit) pH value Oxygen as O2 max. (mg/lit) Fe+Cu+Ni max. (mg/lit) Total solids, alkalinity Silica Oil Organic matter
Water Tube Boiler 20 bar 10 0.05 ----(1) -----
40 bar 60 bar 2 0.5 8.5 – 9.5 0.02 0.01 --0.02 ----(1) (1) ---------
Fire Tube Boiler
View more...
Comments
